Subcontract Report

NREL/SR-560-46674

September 2009

Technoeconomic Boundary 

Analysis of Biological Pathways 

to Hydrogen Production 

March 27, 2008 – August 31, 2009

B.D. James, G.N. Baum, J. Perez, and K.N. Baum 

Directed Technologies, Inc. 

Arlington, Virginia
National Renewable Energy Laboratory

1617 Cole Boulevard, Golden, Colorado 80401-3393

303-275-3000 • www.nrel.gov

NREL is a national laboratory of the U.S. Department of Energy

Office of Energy Efficiency and Renewable Energy

Operated by the Alliance for Sustainable Energy, LLC

Contract No. DE-AC36-08-GO28308

Subcontract Report

NREL/SR-560-46674

September 2009

Technoeconomic Boundary 

Analysis of Biological Pathways 

to Hydrogen Production 

March 27, 2008 – August 31, 2009

B.D. James, G.N. Baum, J. Perez, and K.N. Baum 

Directed Technologies, Inc. 

Arlington, Virginia

NREL Technical Monitor: Ali Jalalzadeh-Azar

Prepared under Subcontract No. AFH-8-88601-01
This publication was reproduced from the best available copy

Submitted by the subcontractor and received no editorial review at NREL

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iii

Preface

The work described in this report was performed under contract to the National Renewable 

Energy Lab (NREL), under the technical guidance of members of the Biological Hydrogen 

Working Group. 

The project considered multiple pathways for the biological production of gaseous 

hydrogen, including photobiological H2 production from a variety of genetically engineered 

algae and bacteria, dark fermentation of waste photobiological organisms, dark fermentation 

of lignocellulosic biomass, and the microbial electrolysis of fermentative waste. 

Additionally, the integration of multiple systems for added hydrogen production and 

reduced cost was considered. This study was based on the best possible designs using data 

currently available, and is not inclusive of all potential designs. 

Within the photobiological analysis, five different organisms were examined including: 

• A truncated antennae Chlamydomonas mutant with an oxygen-tolerant hydrogenase 

• A truncated antennae Cyanobacteria mutant with an oxygen-tolerant hydrogenase 

• A sulfate-permease Chlamydomonas mutant with a truncated antennae 

• An immobilized, sulfur-deprived Chlamydomonas mutant with a truncated antennae 

• A truncated antennae Purple Non-Sulfur (PNS) photosynthetic bacterial mutant 

For each organism, hydrogen production characteristics were defined, individual reactors 

conceptually designed, and levelized hydrogen costs calculated. In keeping with a 

technoeconomic boundary analysis, organism performance was based on the authors 

projection of future genetically modified organisms rather than current laboratory 

experimental measurements. Assumptions were carefully stated and compared to current 

demonstrated performance levels to allow the reader to assess the reasonableness of the 

future performance levels. Additionally, the hydrogen cost derived from photobiological 

systems achieving a near term energy efficiency were also computed to assist in bounding 

the expected future price of hydrogen. 

For the fermentative systems, three main systems were examined: 

• H2 production using dark fermentation of photobiological systems algal waste 

• H2 production using dark fermentation of lignocellulosic feedstock (corn stover)

• H2 production from a Microbial Electrolysis Cell (MEC) using fermentation waste as 

a feedstock. 

While MEC systems produce hydrogen via electrolysis rather than fermentation, we group 

them with the fermentation systems for convenience as they are linked via a wasteproduct/feedstock relationship. The design and performance of the MEC plant draws 

heavily from the concepts and laboratory work conducted at Penn State University. The 

design of the lignocellulosic fermentation plant was based largely on a detailed NREL 

report analyzing the performance and cost of ethanol production from corn stover. 
iv

Fermentation for ethanol production and fermentation for hydrogen production share many 

characteristics. Consequently, the current project work product was greatly enhanced by 

making use of the analogous analysis. For each system, a plant design was selected, its 

capital cost and performance were estimated, and resulting levelized hydrogen cost 

computed. The DOE H2A cost estimation spreadsheet tool was used to estimate these 

levelized hydrogen costs so as to allow easy and transparent comparisons to other 

production options. 

Finally, we integrated the most logical combinations of the above enumerated systems based 

on synergies in plant design, seeking to realize hydrogen cost reduction through combined 

utilization of reactor components and potential use of waste products as feedstock. 

Hydrogen cost from the integrated systems is then compared to hydrogen cost from the 

stand-alone systems. 

Project results were reported in four sequentially prepared reports. This project final report 

consists of these four reports plus a comprehensive discussion section, all integrated under a 

common cover. Thus the report consists of five sections:

• Part I: Photobiological H2 Production Systems

• Part II: Algae Fermentative H2 Production Systems 

• Part III: Lignocellulosic Fermentative H2 Production Systems & Microbial 

Electrolysis Cell Systems

• Part IV: Integrated Systems 

• Part V: Discussion of Project Results

Since the individual reports were done sequentially, Part V of the report contains a summary 

of all contract work, including recommendations for future actions. The reader is advised to 

look to Part V for an overall project summary and listing of conclusions. 
v 

Acknowledgements

The authors of this report would like to acknowledge the contributions of the Biological

Hydrogen Working Group, whose members provided invaluable technical and/or 

programmatic guidance throughout the analysis project. 

Working Group members included: 

Dr. Ali Jalalzadeh-Azar, Senior Engineer II and Technical Monitor, NREL

Dr. Maria L. Ghirardi, Principal Scientist, NREL

Ms. Pin-Ching Maness, Principal Scientist, NREL

Dr. Michael Seibert, Research Fellow, NREL 

Dr. Anastasios Melis, Professor, UC Berkeley

Dr. George Sverdrup, Laboratory Program Manager, NREL

Ms. Roxanne Garland, Technology Development Manager, US DOE, EER&E

Additionally, other members of the hydrogen production community contributed to this 

analysis through their expert guidance. They are: 

Dr. G. Charles Dismukes, Professor, Princeton University 

Dr. Bruce E. Logan, Professor, The Pennsylvania State University 

The work described in this report was performed under contract to the National Renewable 

Energy Lab (NREL) with sponsorship provided by the US Department of Energy, Office of 

Fuel Cell Technologies, DOE Office of Energy Efficiency and Renewable Energy.
vi 

Contents 

PREFACE....................................................................................................................................................III

PART I: PHOTOBIOLOGICAL H2 PRODUCTION SYSTEMS .............................................................................. 1

1. BASIC SCIENCE................................................................................................................................... 2

1.1 SOLAR ASSUMPTIONS .................................................................................................................................. 2

1.2 BED DEPTH, ALGAL CONCENTRATION, & H2 PRODUCTION RATE ASSUMPTIONS...................................................... 4

2. PHOTOBIOLOGICAL SYSTEMS - BIOLOGICAL AND SYSTEM PARAMETERS .......................................... 13

2.1 NUTRIENTS .............................................................................................................................................. 17

2.2 O2-TOLERANT HYDROGENASE (B-1 & B-2)....................................................................................................23

2.3 SULFATE-PERMEASE GREEN ALGAE (B-3) ......................................................................................................25

2.4 IMMOBILIZED SULFATE-DEPRIVED GREEN ALGAE (B-4)......................................................................................27

2.5 PURPLE NON-SULFUR BACTERIA (B-5)...........................................................................................................30

3. ENGINEERING PARAMETERS............................................................................................................ 31

3.1 PHOTO BIO REACTOR BED (PRIMARY SYSTEM) ...............................................................................................32

3.2 ORGANISM FEED SUBASSEMBLY...................................................................................................................42

3.3 RECYCLE SUBASSEMBLY ..............................................................................................................................46

3.4 GAS CAPTURE SUBASSEMBLY.......................................................................................................................49

3.5 CONTROL SYSTEM SUBASSEMBLY .................................................................................................................57

4. BILL OF MATERIALS FOR PATHWAYS................................................................................................ 62

4.1 B-1 PATHWAY PRODUCTION PLANT .............................................................................................................62

4.2 B-2 PATHWAY PRODUCTION PLANT .............................................................................................................64

4.3 B-3 PATHWAY PRODUCTION PLANT .............................................................................................................66

4.4 B-4 IMMOBILIZED PHOTOBIOLOGICAL SYSTEM................................................................................................68

4.5 B-5 PATHWAY PRODUCTION PLANT .............................................................................................................71

5. CAPITAL COST ASSUMPTIONS AND CALCULATIONS .......................................................................... 73

5.1 CAPITAL COST ASSUMPTIONS ......................................................................................................................73

5.2 CALCULATIONS EXPLAINED..........................................................................................................................76

6. LEVELIZED COSTS ASSUMPTIONS & CALCULATIONS (H2A)................................................................ 80

6.1 STANDARD H2A PARAMETERS .....................................................................................................................81

6.2 PATHWAY COMMON PARAMETERS...............................................................................................................83

6.3 PATHWAY SPECIFIC PARAMETERS .................................................................................................................84

6.4 RESULTS FOR LEVELIZED HYDROGEN COSTS ....................................................................................................88

6.5 NEAR TERM PERFORMANCE RESULTS............................................................................................................89

7. DISCUSSION OF RESULTS ................................................................................................................. 90

7.1 CAPITAL COSTS ......................................................................................................................................... 90

7.2 FIXED O&M (PLANT LABOR) .......................................................................................................................92

7.3 VARIABLE COSTS ....................................................................................................................................... 93

7.4 FEEDSTOCK COSTS..................................................................................................................................... 93

7.5 REACTOR FOOTPRINT................................................................................................................................. 94

8. PHOTOBIOLOGICAL SYSTEM CONCLUSIONS AND RECOMMENDATIONS ........................................... 94

9. PART 1 APPENDIX A: KEY TERMS AND DEFINITIONS ......................................................................... 96

10. PART 1 APPENDIX B: MASS BALANCE............................................................................................... 98
vii 

PART II: ALGAE FERMENTATIVE H2 PRODUCTION SYSTEMS ...................................................................... 99

11. INTRODUCTION ............................................................................................................................. 100

12. THEORETICAL AND PRACTICAL FERMENTATION REACTIONS........................................................... 101

13. ALGAE FERMENTATION TESTS........................................................................................................ 102

13.1 FERMENTOR FEEDSTOCK ......................................................................................................................104

14. BASIC FERMENTATION SYSTEM DIAGRAM..................................................................................... 105

14.1 FERMENTATION OPERATION SCHEME.....................................................................................................106

14.2 FERMENTOR SUBASSEMBLY ..................................................................................................................107

15. CAPITAL COST OF THE FERMENTATION SYSTEM............................................................................. 107

15.1 TANK COST ....................................................................................................................................... 108

16. FERMENTATION OUTPUTS FOR C-1 THROUGH C-5 ......................................................................... 108

17. LEVELIZED COSTS ASSUMPTIONS & CALCULATIONS (H2A).............................................................. 109

17.1 STANDARD H2A PARAMETERS ..............................................................................................................109

17.2 PATHWAY COMMON PARAMETERS ........................................................................................................111

17.3 PATHWAY SPECIFIC PARAMETERS ..........................................................................................................112

17.4 LEVELIZED COSTS................................................................................................................................ 114

18. COST RESULTS FOR THE FERMENTATIVE SYSTEMS.......................................................................... 114

19. PROCESSING OF FERMENTOR LIQUID OUTPUTS ............................................................................. 115

20. ALGAE FERMENTATION SYSTEM CONCLUSIONS AND RECOMMENDATIONS ................................... 117

PART III: LIGNOCELLULOSIC FERMENTATIVE H2 PRODUCTION SYSTEMS.................................................. 118

& MICROBIAL ELECTROLYSIS SYSTEMS.................................................................................................... 118

21. LIGNOCELLULOSIC HYDROGEN PRODUCTION................................................................................. 119

21.1 FERMENTATION REACTIONS..................................................................................................................120

21.2 ASSUMPTIONS ................................................................................................................................... 121

21.3 LIGNOCELLULOSIC FERMENTATIVE PARAMETERS .......................................................................................122

21.4 ENGINEERING PARAMETERS..................................................................................................................123

21.5 CORN STOVER PREP SUBASSEMBLY ........................................................................................................124

21.6 PRETREATMENT/HYDROLYSIS SUBASSEMBLY............................................................................................127

21.7 FERMENTATION SUBASSEMBLY..............................................................................................................130

21.8 SEED PRODUCTION SUBASSEMBLY .........................................................................................................132

21.9 STORAGE SUBASSEMBLY ......................................................................................................................134

21.10 WASTEWATER TREATMENT SUBASSEMBLY ..............................................................................................136

21.11 GAS COMPRESSION AND SEPARATION SUBASSEMBLY ................................................................................139

21.12 BILL OF MATERIALS.............................................................................................................................141

22. MICROBIAL ELECTROLYIS CELL (MEC) HYDROGEN PRODUCTION..................................................... 144

22.1 THEORETICAL REACTION AND ASSUMPTIONS ...........................................................................................144

22.2 MEC OPERATING PARAMETERS ............................................................................................................144

22.3 MEC SYSTEM DESCRIPTION .................................................................................................................146

22.4 BILL OF MATERIALS.............................................................................................................................156

23. COST ASSUMPTIONS AND CALCULATIONS ..................................................................................... 158

23.1 VARIABLE COSTS ................................................................................................................................ 158

24. HYDROGEN PRODUCTION COSTS FOR FERMENTOR AND MEC........................................................ 163
viii

25. LIGNOCELLULOSE FERMENATION AND MEC SYSTEM CONCLUSIONS AND RECOMMENDATIONS ..... 164

25.1 LIGNOCELLULOSE FERMENTATION SYSTEM ..............................................................................................164

25.2 MICROBIAL ELECTROLYSIS CELL (MEC) SYSTEM .......................................................................................165

26. PART III APPENDIX A: FERMENTOR MASS/HEAT BALANCE ............................................................. 166

27. PART III APPENDIX B: MEC MASS/HEAT BALANCE .......................................................................... 168

PART IV: INTEGRATED SYSTEMS ............................................................................................................. 169

28. SYSTEMS INTEGRATION................................................................................................................. 170

28.1 STACKED PHOTOBIOLOGICAL PATHWAY ..................................................................................................171

28.2 PHOTOBIOLOGICAL-FERMENTOR INTEGRATION ........................................................................................173

28.3 LIGNOCELLULOSIC FERMENTATION/MEC INTEGRATION.............................................................................175

29. RESULTS AND DISCUSSION FOR INTEGRATED SYSTEMS.................................................................. 178

30. INTEGRATED SYSTEM CONCLUSIONS AND RECOMMENDATIONS ................................................... 180

31. PART IV APPENDIX A: INTEGRATED FERMENTOR/MEC MASS/HEAT BALANCE ................................ 181

PART V: DISCUSSION OF PROJECT RESULTS ............................................................................................ 184

32. SUMMARY OF RESULTS AND CONCLUSIONS .................................................................................. 185

32.1 PHOTOSYNTHESIS SYSTEMS...................................................................................................................185

32.2 ALGAE FERMENTATION SYSTEMS ...........................................................................................................186

32.3 LIGNOCELLULOSE FERMENTATION SYSTEM ..............................................................................................187

32.4 MICROBIAL ELECTROLYSIS CELL (MEC) SYSTEM........................................................................................188

32.5 INTEGRATED STACKED PHOTOSYNTHESIS SYSTEM......................................................................................188

32.6 INTEGRATED PHOTOSYNTHESIS AND FERMENTOR SYSTEM ..........................................................................189

32.7 INTEGRATED LIGNOCELLULOSIC FERMENTOR AND MEC .............................................................................189

32.8 GAS PROCESSING ............................................................................................................................... 189

32.9 HYDROGEN PRODUCTION COST COMPARISONS ........................................................................................189

33. RECOMMENDATIONS FOR FUTURE WORK ..................................................................................... 192
ix

Figures

Figure 1-1. Hourly Irradiance .................................................................................................. 3

Figure 1-2. Daily Insolation Variation over a Year ................................................................. 3

Figure 1-3. ETR vs. Actinic Light ........................................................................................... 6

Figure 1-4. Calculation of Peak ETR ....................................................................................... 8

Figure 1-5. Biological Production Limits – detail ................................................................... 9

Figure 1-6. Biological Production Limits – total potential ...................................................... 9

Figure 1-7. Hydrogen Production Efficiency - Wild-Type (WT) .......................................... 10

Figure 1-8. Hydrogen Production Efficiency – Mutant 1 (MT) ............................................ 10

Figure 1-9. Hydrogen Production Efficiency – Mutant 2 (M2T) .......................................... 10

Figure 1-10. Hydrogen Production - Wild-Type (WT) ......................................................... 11

Figure 1-11. Hydrogen Production – Mutant 1 (MT) ............................................................ 11

Figure 1-12. Hydrogen Production – Mutant 2 (M2T) .......................................................... 11

Figure 1-13. Hourly Hydrogen Production Variation – M2T ................................................ 12

Figure 2-1. Biological Parameters of Photobiological Systems ............................................ 14

Figure 2-2. TAP Medium Costs ............................................................................................. 18

Figure 2-3. Fertilizer prices per ton since 2000. .................................................................... 18

Figure 2-4. Nutrient Rates Computed .................................................................................... 19

Figure 2-5. Sample Fertilizer Costs ....................................................................................... 20

Figure 2-6. Acetic acid Consumption by PNS ....................................................................... 21

Figure 2-7. Carbon Dioxide Solubility in Water ................................................................... 23

Figure 2-8. Oxygen-Tolerant Hydrogenase Process .............................................................. 25

Figure 2-9. Sulfate Permease Process .................................................................................... 27

Figure 2-10. Immobilized Sulfur Deprived Process .............................................................. 29

Figure 2-11. Purple Non-Sulfur Bacteria Process.................................................................. 31

Figure 3-1. Simplified Photobiological Hydrogen Production Plant ..................................... 32

Figure 3-2. Strengths and Weaknesses of Single Bed Reactors ............................................ 33

Figure 3-3. Conceptual Design: Single Bed Reactors ........................................................... 33

Figure 3-4. Strengths and Weaknesses of Dual Bed Reactors ............................................... 34

Figure 3-5. Conceptual Design of Dual Bed Reactors ........................................................... 34

Figure 3-6. Strengths and Weaknesses of Chemostat Bed Reactors ..................................... 35

Figure 3-7. Conceptual Design of Chemostat Bed Reactors ................................................. 35

Figure 3-9. Conceptual Design of Chemostat II Bed Reactors .............................................. 36

Figure 3-10. Reactor Bed Design........................................................................................... 38

Figure 3-11. Reactor Bed Components .................................................................................. 38

Figure 3-12. Polyethylene Hydrogen Losses ......................................................................... 39

Figure 3-13. Film-Frame interface point ............................................................................... 40

Figure 3-14. Paddlewheel Electrical Requirements ............................................................... 40

Figure 3-15. Capital Costs of Reactor Bed Components ....................................................... 42

Figure 3-16. Organism Feed Components ............................................................................. 43

Figure 3-17. Organism Feed Subassembly Design ................................................................ 43

Figure 3-18. Capital Cost of Organism Feed Components .................................................... 45

Figure 3-19. Recycle Components ......................................................................................... 46

Figure 3-20. Recycle Subassembly Design ........................................................................... 46
x 

Figure 3-21. Types of Centrifuges ......................................................................................... 47

Figure 3-22. Drum Filter Diagram ......................................................................................... 48

Figure 3-23. Capital Costs of Recycle Components .............................................................. 49

Figure 3-24. Gas Capture Components .................................................................................. 49

Figure 3-25. Gas Capture Subassembly Design .................................................................... 50

Figure 3-26. Possible PSA Parameters .................................................................................. 53

Figure 3-27. Nano-porous Membranes .................................................................................. 55

Figure 3-28. Capital Cost of Gas Capture Components ........................................................ 56

Figure 3-29. Control System Components............................................................................. 58

Figure 3-30. Control System Subassembly Design ............................................................... 59

Figure 3-31. Control System Wiring and Conduit Quantities for B-1, B-2 ........................... 60

Figure 3-32. Capital Costs of Control System Components .................................................. 60

Figure 4-1. Production Plant Design for Oxygen-Tolerant Hydrogenase (Chlamy) ............. 63

Figure 4-2. Bill of Materials for B-1 Pathway ...................................................................... 64

Figure 4-3. Production Plant Design for Oxygen-Tolerant Hydrogenase (Cyanobacteria) ... 65

Figure 4-4. Bill of Materials for B-2 Pathway ....................................................................... 66

Figure 4-5. Production Plant Design for Sulfate Permease (Chlamy) ................................... 67

Figure 4-6. Bill of Materials for B-3 Pathway ....................................................................... 68

Figure 4-7. Production Plant Design for Immobilized Sulfur Deprived ................................ 70

Figure 4-8. Bill of Materials for B-4 Pathway ....................................................................... 71

Figure 4-9. Production Plant Design for Purple Non-Sulfur ................................................. 72

Figure 4-10. Bill of Materials for B-5 Pathway ..................................................................... 73

Figure 5-1. Piping sizes for B-1 through B-5 ........................................................................ 75

Figure 5-2. Electricity Consumption per Pathway ................................................................. 78

Figure 5-3. Davis-Bacon Wage Determinations and Blue Book Rental Costs ...................... 79

Figure 5-4. Excavation Cost Estimate for B-1 System using California Costs ..................... 80

Figure 6-1. H2A Default Values and Assumptions used for all Biohydrogen Pathways ...... 82

Figure 6-2. Parameters Common to all Pathways .................................................................. 83

Figure 6-3. Pathway Specific Parameters .............................................................................. 84

Figure 6-4. Land Required for each Pathway ........................................................................ 85

Figure 6-5. Plant Staff Requirements for 1 tonne H2/day plant ............................................. 86

Figure 6-6. Utilities Usage ..................................................................................................... 86

Figure 6-7. Feedstock Usage .................................................................................................. 87

Figure 6-8. LDPE Film Replacement Costs .......................................................................... 87

Figure 6-9. Levelized costs for B-1 through B-5 Pathways ................................................... 88

Figure 6-10. Comparison of Levelized Cost Components..................................................... 89

Figure 6-11. Near Term STH Efficiency Assumptions ......................................................... 89

Figure 6-12. Levelized H2 Costs based on Near Term STH Efficiency ................................ 90

Figure 6-13. Comparison of Levelized Cost Components for Near Term Efficiencies ........ 90

Figure 7-1. Capital Cost Increases from Increased Plant Size ............................................... 91

Figure 7-2. Labor Assumptions for Different Plant Sizes ..................................................... 92

Figure 7-3. Labor Cost Reductions from Increased Plant Size .............................................. 93

Figure 7-4. Normalized Hydrogen Production ...................................................................... 94

Figure 8-1. Capital Cost Allocations by Subassembly .......................................................... 95

Figure 11-1. Fermentation Process within a Biological System .......................................... 101

Figure 12-1. Potential Fermentative Feedstock Materials ................................................... 102

Figure 13-1. NREL Fermentation Experimental Data ........................................................ 103
xi 

Figure 14-1. Fermentation System Design .......................................................................... 106

Figure 15-1: BOM for Fermentor Systems C-1 and C-2 (sized for the organism wastestream of a 10TPD B-1 or B-2 Photobiological Hydrogen production system) ...... 108

Figure 15-2: BOM for 10 TPD Fermentor System C-5 (sized for the PNS waste-stream of a 

10TPD B-5 Photobiological Hydrogen production system) .................................... 108

Figure 16-1. Fermentative Hydrogen Ouput (from 10TPD Photobiological System) ......... 109

Figure 17-1. H2A Default Values and Assumptions used for all Fermentative Pathways .. 110

Figure 17-2. Parameters Common to all Pathways .............................................................. 111

Figure 17-3. Pathway Specific Parameters .......................................................................... 112

Figure 17-4. Land Required for each Pathway .................................................................... 113

Figure 17-5. Plant Staff Requirements for Fermentor plant ................................................ 113

Figure 17-6. Electricity Usage ............................................................................................. 113

Figure 18-1: H2A Model Projected Hydrogen Cost per kg from Fermentor Systems C-1, 

C-2, and C-5 ............................................................................................................. 114

Figure 19-1: Fermentor Outputs .......................................................................................... 115

Figure 19-2: MEC Flows ..................................................................................................... 116

Figure 19-3. C-1/C-2 Fermentation + MEC ........................................................................ 116

Figure 19-4. C-5 Fermentation ............................................................................................ 117

Figure 21-1. Potential Fermentative Feedstock Materials ................................................... 119

Figure 21-2. Lignocellulosic Biomass ................................................................................. 119

Figure 21-3. Parameters for Fermentation System .............................................................. 122

Figure 21-4. Simplified Lignocellulosic Hydrogen Production Plant ................................. 124

Figure 21-5. Corn Stover Prep Components ........................................................................ 124

Figure 21-6. Corn Stover Prep Subassembly Design ........................................................... 126

Figure 21-7. Capital Cost of Corn Stover Prep Subassembly .............................................. 127

Figure 21-8. Pretreatment/Hydrolysis Components ............................................................ 128

Figure 21-9. Pretreatment/Hydrolysis Subassembly Design ............................................... 129

Figure 21-11. Fermentation Components ............................................................................ 130

Figure 21-12. Fermentation Subassembly Design ............................................................... 131

Figure 21-13. Capital Costs of Fermentation Subassembly ................................................. 132

Figure 21-14. Seed Production Components ....................................................................... 132

Figure 21-15. Seed Production Subassembly Design .......................................................... 133

Figure 21-16. Capital Costs of Seed Production Subassembly ........................................... 134

Figure 21-18. Storage Subassembly Design ........................................................................ 135

Figure 21-19. Capital Costs of Storage Subassembly .......................................................... 136

Figure 21-20. Wastewater Treatment Subassembly Components ....................................... 137

Figure 21-21. Wastewater Treatment Subassembly Design ................................................ 138

Figure 21-22. Capital Cost of Wastewater Treatment Subassembly ................................... 139

Figure 21-23. Gas Compression and Separation Components ............................................ 140

Figure 21-24. Gas Compression and Separation Subassembly Design ............................... 141

Figure 21-25. Capital Cost of Gas Compression and Separation Subassembly .................. 141

Figure 21-26. Lignocellulosic Fermentation Flow Chart .................................................... 142

Figure 21-27. Bill of Materials Lignocellulosic Fermentation ........................................... 143

Figure 22-1. Parameters for MEC System ........................................................................... 145

Figure 22-2. Simplified MEC Hydrogen Production Unit ................................................... 146

Figure 22-3. MEC Unit Field Layout .................................................................................. 147

Figure 22-4. MEC Components ........................................................................................... 148
xii 

Figure 22-5. MEC Cell ........................................................................................................ 148

Figure 22-6: MEC Operating Performance – Pt Cathode .................................................... 149

Figure 22-7. MEC Operating Performance - Brush Cathode .............................................. 150

Figure 22-8. MEC Cost Optimization .................................................................................. 151

Figure 22-9: MEC Test Unit Top view with Half-Brush Cathode and Anode ................... 152

Figure 22-10. Capital Costs of MEC Subassembly ............................................................. 153

Figure 22-11. Storage Components ..................................................................................... 153

Figure 22-12. Storage Subassembly Design ........................................................................ 154

Figure 22-13. Capital Costs of Storage Subassembly .......................................................... 154

Figure 22-14. Gas Compression and Separation Components ............................................ 155

Figure 22-15. Gas Compression and Separation Subassembly Design ............................... 156

Figure 22-16. Capital Cost of Gas Compression and Separation Subassembly .................. 156

Figure 22-17. MEC System Flow Chart .............................................................................. 157

Figure 22-18. MEC Stand-Alone Bill of Materials ............................................................. 157

Figure 23-2. Power Requirements MEC System ................................................................. 160

Figure 23-3. Labor Breakdown - Fermentor ........................................................................ 160

Figure 23-4. Additional Consumables Costs – Fermentation system .................................. 161

Figure 23-5. Water Use ........................................................................................................ 162

Figure 23-6. Total Variable Costs for Fermentor ................................................................ 162

Figure 23-7. Total Variable Costs for MEC ........................................................................ 162

Figure 24-1: Total Hydrogen Cost ....................................................................................... 163

Figure 24-2. Comparison of Levelized Cost Components................................................... 163

Figure 24-3. Fermentor Liquid-Organic Output Value........................................................ 164

Figure 28-1. Potential Biological Integration Methods ....................................................... 170

Figure 28-2. Process Diagram – Integrated Photobiological Systems ................................. 172

Figure 28-3. Tiered Reactor Bed Conceptual Design .......................................................... 172

Figure 28-4. Bill of Materials for B-3/B-5 Stacked Beds .................................................... 173

Figure 28-5. Photobio-Fermentor Process Diagram ............................................................ 174

Figure 28-6. BOM for B-5/C-5 Integrated System .............................................................. 175

Figure 28-7. Process Diagram for Fermentation-MEC Integration ..................................... 176

Figure 28-8. Bill of Materials for Integrated Fermentation-MEC ....................................... 177

Figure 28-9. Fermentor/MEC Mass Balance Flow Chart .................................................... 178

Figure 29-2. Comparison of Integrated System Costs ......................................................... 179

Figure 32-1. Photobiological System STH Energy Efficiency ............................................ 186

Figure 32-2. H2 Production Costs ........................................................................................ 190
1 

Part I: Photobiological H2 Production 

Systems 
2 

Introduction 

Directed Technologies Inc. (DTI) is under contract to the National Renewable Energy Lab 

(NREL) to conduct a techno-economic evaluation of photobiological hydrogen production 

systems. This report documents the biological and engineering characteristics of five algal 

and bacterial hydrogen production systems selected by DOE and NREL for evaluation. 

Those characteristics are divided into three areas, namely, basic science, biology, and 

engineering. Since basic science assumptions are strongly correlated with biological 

parameters, they are provided first. Hydrogen production capabilities are a function of the 

reactor bed design so those details are provided in the engineering section of this report. 

Because product hydrogen product is combined with other gases, the characteristics of a gas 

separation system will be given additional attention. The last section of this Task B report 

utilizes the information collected to estimate a final result, given as a cost per unit of 

produced hydrogen. 

1.

1.1 Solar Assumptions

Basic Science 

Solar insolation is a key factor in determining algal growth and H2 production rates. 

Selecting a geographical location provides ambient temperatures, hours of exposure, and 

amount of irradiance. For the purpose of this study we have based our solar input 

assumptions on the solar irradiance properties of the Phoenix, AZ area. Hourly solar 

irradiance (W/m2 striking the horizontal reactor bed) for this area is shown in Figure 1-1. 

The NREL SOLPOS model is used for solar position angles and intensity1. This model 

calculates extra terrestrial radiation (ETR) by the hour and day throughout the year. The 

total clearness index, Kt, is multiplied by the ETR to obtain the terrestrial irradiance on the 

horizontal surface. Kt values for the relevant months in the Phoenix area are obtained from 

the NASA Atmospheric Data Center / Surface meteorology and Solar Energy (SSE) 

database2. The 3 curves plotted in Figure 1-1 show the solar peak at June 21, the autumnal 

equinox at September 21, and the solar minimum at December 21. These curves also show 

the seasonal peak radiation which determines the peak potential H2 production for those 

periods. The peak June irradiance at 964W/m2 is slightly below the nominal peak value of 

1,000 W/m2

 due to the sun’s maximum elevation angle being 80 degrees above the horizon. 

 

1 NREL MIDS SOLPOS (Solar Position and Intensity) model may be found at 

http://rredc.nrel.gov/solar/codesandalgorithms/solpos/. 2 NASA Atmospheric Data Center, Langley ASDC User Services, Surface meteorology and Solar Energy 

(SSE) data base (release 6.0).
3 

Figure 1-1. Hourly Irradiance

Hourly Irradiance

Phoenix (33N,112 W), Horizontal Plane 

0

100

200

300

400

500

600

700

800

900

1000

4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

Solar time (hours after midnight)

Irradiance W / m2

Jun 21, solar peak

Sep 21

Dec 21, solar low

NREL SOLPOS - ETR 

 + max clearness index

The daily insolation energy (kW-hr/m2

/day) on a horizontal surface throughout the year for 

the Phoenix area was obtained from the NASA SSE database, and is plotted in Figure 1-2. 

These plotted values are averages for the respective months. The yearly mean value of 

5.5kW-hr/m2/day is also shown. 

Figure 1-2. Daily Insolation Variation over a Year 

Monthly Insolation 

Phoenix (33N,112W), Horizontal Plane

0

1

2

3

4

5

6

7

8

9

Jan Feb Mar Apr May Jun Jul Aug Sep Oct Nov Dec

 Insolation kW-hr / m2 / day

average over month

yearly mean

NASA Atmospheric Science Data Center
4 

Note that the effects of clouds are implicitly included in these calculations, as the monthly 

data points are for the average day including the average amount of clouds. The normalized 

cell growth and hydrogen production parameters used in the following analyses have 

assumed an average sunlight of 5.5kW-hr/m2/day to generate a usable hydrogen production 

rate of 1,000kg/day averaged over the year. 

1.2 Bed Depth, Algal Concentration, & H2 Production Rate Assumptions

The depth of the reactor bed and algal concentration are matched to ensure full photon 

capture by the organisms. If organism concentration is too dilute for a given total depth, 

light will penetrate to the bottom of the bed and photons will be “wasted”. Should the 

organism be too concentrated, photons will be fully absorbed in only the upper layers of the 

bed potentially “starving” the organisms at the lower depths of photons and adversely 

affecting their health3. The organism concentration and bed depth are also important 

because they impact specifications for other system components such as pumps and valves, 

type of solids separation equipment, and type of mixing equipment. 

Choice of bed depth and concentration depends on the mode of growth, batch versus 

continuous dilution. For batch mode, the ideal fixed-volume bioreactor to use for achieving 

constant light transmission throughout a growing population of dispersed cells would have 

to allow for continuous decrease in path length as the population grows to higher cell 

densities. Alternatively, a continuous dilution bioreactor of fixed volume could operate with 

a fixed cell density after reaching steady-state growth and thus would have constant 

absorbance. If the latter growth mode is not used and replaced with batch culture growth 

mode then the bed depth must be chosen so that the absorbance after reaching stationary 

phase, e.g., at beginning of the photo-hydrogen production stage, produces the maximum 

rate of hydrogen production. 

1.2.1 Beers Law Photon Absorption Model

The rate of photon absorption and thus the maximum hydrogen production at a given bed 

depth increment can be determined by Beer's Law. At each depth increment, Δd, one 

calculates the number of photons absorbed and then multiplies that by the specific rate of H2

production in molecules per photon (this is equivalent to the light saturation curve for photohydrogen production which is an experimental measurement distinct for each cell type or 

mutant).

Beer’s law states that:

A = b*C*ε where A= Absorbance (unitless) 

 b= depth into liquid in cm 

 C= cell concentration in gdrywt/L

 ε= absorption coefficient in L/cm/gdrywt

Absorbance is defined as the negative logarithm of the ratio of exiting light intensity to 

entering (incident) light intensity: 

A = -log(I/Io) 

 

3 Proper bed mixing can mitigate or potentially eliminate the effects of shallow light penetration.
5 

From this definition, I/Io (light intensity I, at a given bed depth, b) is derived from: 

 

I/Io = 10-εbC

To determine the required bed depth and cell concentration to maximize photon capture by 

the organisms, we integrate the above equation over the depth of the bed. Note that for this 

approach and Beer’s law to be valid, we need to make multiple assumptions:

1. The cell volume and pigment content per cell must not change appreciably during 

the lifecycle of the sample. While most cells increase pigment content and volume as 

they age and, therefore, do not strictly obey Beer’s law, we assume average rather 

than instantaneous cellular characteristics for purposes of the calculation. As shall 

be discussed later in the report, this assumption is most valid for the Chemostat II 

system, where a constant cell concentration is maintained.

2. The cells must not be allowed to aggregate, as commonly occurs without adequate 

mixing, resulting in biofilm formation. Consequently, this analysis is most 

applicable to planktonic (or free-floating) cells. With the exception of the B-4 

System (see p. 25), all systems considered are free-floating cells in an actively mixed 

medium. This analysis is not relevant to the B-4 immobilized system.

3. The cell density must be sufficiently low that cell “shading” does not occur. In the 

present configuration shading is minimized by dilution: all systems, with the 

exception of B-3 and B-4 for which this analysis does not apply due to the necessity 

of higher cell concentration as a feedstock for oxidative respiration, maintain a cell 

concentration of less than 0.2gdrywt/L.

4. Corrections for the light lost from the bioreactor arising from light scattering of cells 

(turbidity) are generally estimated to be low of order: 10% for the cell concentrations 

considered in this report and for a normal incident angle. Generally, visible light 

scattering by cells of the size of Chlamydomonas (10 microns) is predominantly 

forward scattering and thus most of the scattered light is retained in the bioreactor. 

Thus, no correction for light scattering is included in the analysis. 

1.2.2 Light Saturation

While Beer's Law relates photon absorption to cell and bed characteristics, there is an 

additional limitation on H2 production due to saturation of photon/electron conversion 

capability within the cells. There is a charge transfer limitation (alternately called a 

saturation limit) within the photosynthetic pathway that limits the rate at which 

photons/electrons can be processed by the PSII reaction. This consequently limits 

utilization of incident photons, limiting the rate of the PSI reaction and H2 production, and 

thus lowering the Solar to Hydrogen (STH) energy conversion efficiency. At moderate-tohigh light intensity, this saturation limit is greatly exceeded by wild type organisms and, to a 

lesser extent, by current truncated antenna mutants. Two analyses of the saturation limit 

effects are discussed in this and the following section. 
6 

To estimate this charge transfer limitation, we examine the measured quantum efficiency for 

electron transport by PSII as a function of Actinic Light Photon Flux Density (PFD) for 

whole cells of the green alga Chlorella. Figure 1-3 displays electron transfer rate (ETR) vs. 

actinic light for two methods of ETR calculation4. Actinic light generally refers to portion 

of light that causes a chemical change. In this specific case, light refers to the light within 

the PAR range. These data come from experiments and are qualitatively representative of 

data in the literature for other algae including Chlamydomonas reinhardtii. 

Figure 1-3. ETR vs. Actinic Light 

Based on the theory of light absorption and following well-known Poisson statistics5

 

4 Unpublished data from Gennady Ananeyev, Princeton University.

we 

postulate that the relationship between ETR and light intensity will follow the relationship, 

5

 “Light Saturation Curves and Quantum Yields in Reaction Centers from Photosynthetic Bacteria”, 

Blankenship et al. Journal of Biophysics, Volume 45 February 1984 455-461.
7 

Y/Ymax = 1-e-σΦE

where;

Ymax = Incident light intensity 

Y = Absorbed light intensity 

σ = optical cross section in cm

2 

Φ = quantum yield, probability of charge separation/photon 

E = Incident intensity in photons/cm2/sec

We next analytically determine the σΦ value (σΦ = 0.053346) that results in a curve fit to 

the experimental data. As shown in Figure 1-4, a reasonably good curve fit was obtained. 

While this form of the data and analytic model indicate that ETR peaks at a light intensity of 

approximately 60W/m2, it does not show the peak ETR. Consequently, to determine the 

peak ETR, we postulate that there are no charge transfer limitations at very low light levels 

and thus an extrapolation of the linear portion of the curve will yield an indication of the 

maximum electron transfer. By extrapolating the linear portion of the curve to its 

intersection with the maximum electron transfer (Y/Ymax=1), we estimate that the cells 

would be able to process photons at up to 19W/m2 light intensity if they were unencumbered 

by rate limits. This allows us to compute the peak charge separation rate. Thus we now 

have a measure of both the peak H2 production rate and shape of the light intensity vs. H2

production rate. The light saturation curve can be used together with Beer's law to 

determine the rate of charge separation that occurs at every point within the bioreactor. This 

same approach can be used to calculate the H2 production rate which is assumed to be a 

fixed fraction of the charge separation rate. Thereby the H2 production rate can be obtained 

at each depth increment within the bioreactor and the total absorption summed from all such 

depth increments throughout the bioreactor to obtain the overall H2 production throughout 

the reactor. The result is a substantial decrease in H2 production rate due to light saturation 

effects of typical current organisms. 
8 

Figure 1-4. Calculation of Peak ETR 

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

1.1

0 20 40 60 80 100

Ananyev Data #1

DTI curve fit

Ananyev Data #2

Y/Ymax

Light Intensity (W/m2)

1.2.3 Alternative Model for Photon Absorption Saturation

An extensive analysis of saturation effects on hydrogen production was carried out for 

NREL by Wade Amos in NREL/MP 560-355396. In this report, the hydrogen production 

rate in successive 0.2 mm layers of algae was predicted as a function of depth in the pond, 

the solar intensity, and the algae antenna saturation characteristics. For reduced antenna 

mutants, the analysis predicted major reduction in hydrogen production rate at depths 

shallower than 4 cm due to photon saturation. The photon saturation criterion assumed by 

Amos was that the maximum electron current able to be transferred to an external electron 

acceptor is 1 mole of electrons per gram (dry cell weight) of cells per day. 

As an expansion of the Amos saturation analyses, we investigated the saturation effects 

using a hypothetical mutant (M2T) that had average absorption coefficient (ε) reduced by a 

factor of 5 relative to wild-type (WT). Figure 1-5, based on the Figure 15 of NREL/MP 

560-35593, plots the rate of hydrogen production per m2 per second for each 0.2mm 

increment of bed depth, for three varieties of algae. The thin green curve represents the rate 

of hydrogen production per depth layer of wild-type Chlamydomonas reinhardtii which has 

an average ε of 5.0, if all absorbed photons were successfully converted to hydrogen. The 

thick black line plots the analogous production curve for the Amos report Mutant-Type 

(MT) Chlamydomonas reinhardtii with a reduced antenna, which has an average ε of 2.07. 

The blue line plots the analogous production curve for a hypothetical improved Mutant-2-

Type (M2T) Chlamydomonas reinhardtii with an average ε of 1.0. Figure 1-6 shows the 

 

6 Amos, Wade, “Updated Cost Analysis of Photobiological Hydrogen Production from Chlamydomonas 

reinhardtii Green algae”, NREL/MP 560-35593, January 2004 at page 7 (Gosset, 1995).
9 

full extent of these curves. If all of the PAR photons were absorbed and converted to H2 at 

the rate of 4 photons/one H2 molecule, the hydrogen production would equal the integration 

of the area under each of these absorption curves and the STH energy efficiency would be 

12.2% (assuming the Lower Heating Value (LHV) of H2 = 33.33 kWh/kg).

Figure 1-5. Biological Production Limits – detail

PAR Photon conversion to H2

0

1

2

3

4

5

6

7

8

9

10

0.0 1.0 2.0 3.0 4.0 5.0 6.0 7.0 8.0

total depth - cm

Mole H2/m2/sec per 0.2mm layer X 10^7

100% conversion WT

100% conversion MT

100% conversion M2T 

algae saturation

Figure 1-6. Biological Production Limits – total potential

PAR Photon conversion to H2-Overall 

0

50

100

150

200

250

0.0 1.0 2.0 3.0 4.0 5.0 6.0 7.0 8.0

total depth - cm

Mole H2/m2/sec per 0.2mm layer X 10^7

100% conversion WT

100% conversion MT

100% conversion M2T 

algae saturation

The thick green horizontal line on these curves denotes the algae saturation limit for

hydrogen production based on the epsilon assumptions discussed above. Thus, with 

increasing depth, the actual H2 production rate first follows the saturation thick green line 

for light intensities exceeding the saturation threshold and then follows the appropriate 

absorption curve predicted by Beer’s Law. The ratio of the area below each set of two 

limiting curves to the solar photon flux represents the actual fraction of photons captured for 

H2 production, and thus the process efficiency. 
10 

Shown in Figure 1-7, Figure 1-8 and Figure 1-9 are curves of the H2 production efficiency 

for the wild-type and two mutant types as solar intensity is increased up to the max intensity 

of 1000W/m2

. As can be seen, 100% photon utilization only takes place at low output for 

WT and at levels below 21 W/m2 for MT and 45 W/m2 for M2T. The higher efficiency of 

the Mutant-2-Type (M2T) Chlamydomonas reinhardtii results from the increased photon 

utilization throughout the pond depth. 

Figure 1-7. Hydrogen Production Efficiency - Wild-Type (WT)

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0 100 200 300 400 500 600 700 800 900 1000

Photon Utilization for H2

solar irradiance W/m2

H2 production - WT

Density = 0.2 g/L, 

Algae saturation =2.3 X 10-7 mol H2/m2/s (0.2 mm layer) 

Figure 1-8. Hydrogen Production Efficiency – Mutant 1 (MT)

H2 production - MT

Density = 0.2 g/L, 

Algae saturation =2.3 X 10-7 mol H2/m2/s (0.2 mm layer) 

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0 100 200 300 400 500 600 700 800 900 1000

solar irradiance W/m2

Photon Utilization for H2

Figure 1-9. Hydrogen Production Efficiency – Mutant 2 (M2T)

H2 production - M2T

Density = 0.2 g/L, 

Algae saturation =2.3 X 10-7 mol H2/m2/s (0.2 mm layer) 

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0 100 200 300 400 500 600 700 800 900 1000

solar irradiance W/m2

Photon Utilization for H2
11 

Figure 1-10, Figure 1-11 and Figure 1-12 show resultant hydrogen production rate for both 

wild-type and mutant types as solar irradiance is increased, with the early linear part of the 

curve reflecting full conversion. At full conversion intensities, the saturation limit is not 

exceeded and photon utilization can be 100% (i.e. 100% of the photons absorbed are able to 

be processed by the PSII and PSI reactions.) Output levels off as irradiance increases, with 

the Mutant-2-Type having significantly higher overall production levels. 

Figure 1-10. Hydrogen Production - Wild-Type (WT)

0.0

0.1

0.2

0.3

0.4

0.5

0 100 200 300 400 500 600 700 800 900 1000

g H2/m2/hr

solar irradiance W/m2

H2 production - WT

Density =0.2 g/L, 

Algae saturation =2.3 X 10-7 mol H2/m2/s (0.2 mm layer) 

Figure 1-11. Hydrogen Production – Mutant 1 (MT)

H2 production - MT

Density =0.2 g/L, 

Algae saturation =2.3 X 10-7 mol H2/m2/s (0.2 mm layer) 

0.0

0.1

0.2

0.3

0.4

0.5

0 100 200 300 400 500 600 700 800 900 1000

solar irradiance W/m2

g H2/m2/hr

Figure 1-12. Hydrogen Production – Mutant 2 (M2T)

H2 production - M2T

Density =0.2 g/L, 

Algae saturation =2.3 X 10-7 mol H2/m2/s (0.2 mm layer) 

0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0 100 200 300 400 500 600 700 800 900 1000

solar irradiance W/m2

g H2/m2/hr
12 

To illustrate the combined effects of cell saturation and seasonal and diurnal variation in 

light intensity, hourly hydrogen production curves were plotted in Figure 1-13 for June 21, 

September 21, and December 21. Integrating under the curve gives total daily H2

production and allows an average STH energy conversion efficiency to be calculated. These 

values are shown in Figure 1-13. 

Figure 1-13. Hourly Hydrogen Production Variation – M2T 

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

H2 Production Rate: g H2/m2/hr

Solar time (hours after midnight)

Hourly H2 Production

Mutant2 (ε=1.0), Saturation level: 1 mole electrons/g/day, Density: 0.2 g/L

Phoenix (33N,112 W), Horizontal Plane 

Jun 21, solar peak

Sep 21

Dec 21, solar low

NREL SOLPOS - ETR 

+ max clearness index

Jun 21: 8.66 gH2/m2 (2.6%)

Sep 21: 7.06 gH2/m2 (2.9%)

Dec 21: 5.17 gH2/m2 (4.0%)

In summary, the above curves show H2 production rate for a hypothetical M2T mutant 

antenna algae strain at the following conditions: 

• An organism that obeys Beer’s law 

• Chlamydomonas reinhardtii at cell concentration = 0.2 g/L and a bed depth = 10 cm 

• Mass absorption coefficient ε = 1 L/cm/gdrywt representing a 5:1 reduced antenna 

mutant relative to wild type

• Maximum charge transfer saturation limitation of 1 mole electrons per gram cell dry 

weight of cells per day 

• Linear dependence of the H2 production rate on light intensities below those needed 

to produce electron transfer at the maximum charge transfer limit

• Solar diurnal and yearly fluctuations as described in Section 2.1 

• All photons absorbed by the cell are assumed to impact the chlorophyll antenna 

• All photons successfully entering the PSII/PSI chain productively produce H2 (i.e., 

no photons go toward growth, heat or fluorescence) 

Integrating over the course of the year, we estimate that the resulting annual average STH 

energy conversion efficiency is 3.1%. 

1.2.4 Absorption Parameters Used in This Study

The maximum STH efficiency is highly dependent on the saturation limit assumption. 

While the limitation mechanism is well understood, the upper saturation limit for future 

optimized organisms is poorly defined. The saturation model discussed in the prior sections 

is based on current organism properties. However, this study is focused on potential future 
13 

production using developmental mutants of algae and bacteria which will not have the 

stringent saturation limits of current algae and bacteria. Consequently for this analysis the 

saturation limit was not imposed for the following reasons: 

1. We wish to parametrically assess the cost of hydrogen resulting from multiple 

organism systems and configuration. Consequently, we want to establish a true 

“upper bound” efficiency to gauge the potential of biohydrogen. 

2. We are interested in modeling hypothetical mutant strains that have optimized 

hydrogen generation capability and their exact production characteristics are not 

known. 

3. Mutation of the chlorophyll antenna size is expected to be one approach to 

optimizing STH efficiency and the limits of antenna truncation are not known. 

4. The removal of the saturation limit leads the STH conversion efficiency to be 

constant with respect to light intensity. This in turn means the STH efficiency is 

constant throughout the year and greatly simplifies the reactor bed sizing 

analysis (because it allows us to merely sum total incident photons rather than 

carefully assessing the incident photon variations over the course of the day and 

year).

Assumed STH conversion efficiencies for each of the examined biological systems are 

detailed in the next section. 

2.

“Certain algae and cyanobacteria photoproduce hydrogen for short times as a way to get rid 

of excess energy.”

Photobiological Systems - Biological and System Parameters

7

 Recognizing the potential of using these biological systems as a source 

of energy, scientists have sought out ways to extend and control the duration of hydrogen 

production from these organisms. 

Our analysis focuses on the hydrogen production potential of 5 photobiological systems. 

These are;

1. green algae that co-produces H2 and O2 gases and possesses an O2-tolerant hydrogenase, 

2. cyanobacteria that co-produces H2 and O2 gases and possesses an O2-tolerant 

hydrogenase8

3. sulfate-permease algal mutants that produce only H2 gas due to the effects of a mutated 

sulfate permease gene on the chloroplast, 

, 

4. immobilized sulfate-deprived green algae that produce H2 for extended periods of time, 

and 

5. purple non-sulfur (PNS) bacteria that produce H2 gas in light. 

The baseline conditions that have been used for each of the pathways are summarized in 

Figure 2-1. The Reactor Bed Mode of Operation refers to the process that is used to grow 

algae and produce hydrogen. This mode of operation results in different solar to hydrogen 

(STH) efficiencies for a given organism. The Reactor Bed Mode of Operation is described 

in more detail in the Engineering Parameters section of this report.

 

7 Photobiological Production of Hydrogen (Fact Sheet). FS-560-42285. NREL. Colorado. November 2007.

http://www.nrel.gov/docs/fy08osti/42285.pdf. 8 This strain of cyanobacteria does not currently exist and needs to be genetically engineered.
14 

Figure 2-1. Biological Parameters of Photobiological Systems 

 B-1:

Algae 

O2-tolerant 

Hydrogenase

B-2:

Cyanobacteria

O2-tolerant 

Hydrogenase

B-3:

Algae 

Sulfate 

Permease 

B-4:

Immobilized 

Algae, Sulfur 

deprived 

B-5:

PNS Bacteria

Organism9 C. reinhardtii Synechocystis 

Hydrogenase10

C. reinhardtii C. reinhardtii 

cc124 

Rhodobacter 

sphaeroides 

RV11

Antennae Type LHC (Light 

Harvesting 

Complex) 

deletion Mutant

Phycobilin deletion 

Mutant

LHC deletion 

Mutant

LHC deletion 

Mutant

LHC - II deletion 

Mutant

Reactor Bed Mode of 

Operation.

Chemostat II12 Chemostat II Single-Bed Dual-Bed, 

Single-Bed13

Chemostat II

Cell Growth Conditions

Water Fresh Fresh Fresh Fresh Fresh

Temperature (oC) 25-35 25-35 25-35 25-35 25-35 

Final Concentration 

in solution14

0.2 

 (g/L – Dry Wt)

0.2 

(g/L – Dry Wt)

4.9115

(g/L – Dry Wt)

2.8 

(g/L – Dry Wt)

0.2 

(g/L – Dry Wt)

Inorganic nutrients Fertilizer containing: Potassium, Phosphorous, Nitrogen and trace elements

Daily Rates

g/g organism

K: 3.071 x10-6, P: 4.913 x10-5 , N: 6.142 x10-6

Daily Rates

g/m2

K: 6.14 x10-5 P: 9.83 x10-4 , N: 1.23 x10-4

Organic nutrient None None None None Acetic Acid

Daily Rate

g/g organism N/A N/A N/A N/A 0.624 

Daily Rate

g/m2 N/A N/A N/A N/A 12.5 

Daily CO2 Req16

 g CO2/g dry mass

 0.73 0.73 0.73 0.73 None

Growth duration 2 days initial 4 days 1 day 2 days initial

 

9 We assume an engineered antenna for all organisms. B-1 and B-2 also possess an O2 tolerant hydrogenase, 

thus creating a “double mutant”. While without precedent in the literature, multiple mutations are necessary to 

achieve high STH conversion efficiencies.

10 Based on Synechocystis PCC6803 mutant work of Pin-Ching Maness (NREL)

11 Analysis does not rely on a specific strain of PNS proteobacterium. Multiple examples of possible 

organisms are cited in “Photobiological hydrogen production: photochemical efficiency and bioreactor 

design,” Akkerman et. al., Intl. Journal of Hydrogen Energy 27 (2002) 1195-1208.

12 Chemostat II is a chemostat with simultaneous growth and H2 production, described in Section 4.1.4 

13 As discussed in Section 3.4, this system utilizes a two day production period, followed by a two day growth 

period in a single bed. At the end of 180 days of production, the algae grown in a second bed is transferred to 

the production bed, replacing the spent algae. In this way, it is both a dual-bed and single-bed system 14 Peak cell concentration in growth bed at end of growth phase, and the target starting concentration for H2

production. For the Chemostat II systems (B-1, B-2, and B-5), this concentration refers to the final 

concentration at the end of the initial grown phase i.e. the growth phase used to create the initial colony prior 

to chemostat operation.

15 Concentrations higher than 0.2 are needed for B3 and B4 in order to ensure that enough cell mass is 

accumulated for respiration to keep the system anaerobic

16 Four times the sustainment amount used for production conditions in this table. Based on private 

communication with A. Melis. (University of California-Berkeley)
15 

B-1 B-2 B-3 B-4 B-5 

Hydrogen Production Conditions

Water Fresh Fresh Fresh Fresh Fresh

Temperature (oC) 25-35 25-35 25-35 25-35 25-35 

Cell Concentration 

in solution 

0.2 

(g/L – Dry Wt)

0.2 

(g/L – Dry Wt)

0.85 

(g/L – Dry Wt)

1.8117

(g/L – Dry Wt)

0.2 

(g/L – Dry Wt)

Nutrients Fertilizer containing 

Potassium, Phosphorous, and Nitrogen and trace elements

Fertilizer 

containing 

Potassium, and 

Nitrogen and trace 

elements + Low 

Level of Sulfur

Fertilizer 

containing 

Potassium, 

Phosphorous, 

and trace 

elements + 

Acetic Acid

Daily Rates

g/g organism

K: 7.677 x10-7

P: 1.228 x10-5

N: 1.535 x10-6

K: 7.677 x10-7

N: 1.535 x10-6 

S: very low levels

K: 3.071 x10-6

P: 4.913 x10-5

Ac acid: 2.1

Daily Rates

g/m2

K: 1.54 x10-5

P: 2.46E x10-4 

N: 3.07 x10-5

K: 1.54 x10-5

N: 3.07 x10-5

S: very low levels

K: 6.14 x10-5

P: 9.83 x10-4

Ac acid: 42.15 

Daily CO2 Req 

g CO2/g dry mass

0.1618 0.1618 None None None

Oxygen Tolerance Aerobic/O2 

tolerant

Aerobic/O2 tolerant Anaerobic Anaerobic19 Anaerobic

Deprivations from 

media

None None None Phosphate 

deprived, Sulfur 

limited

Nitrogen 

deprived

Production Cycle 

Duration 

Semi-infinite, 

continuous

Semi-infinite, 

continuous

3 days H2

production, 

4 days growth

 3 days H2

production, 1 day 

growth for ~180 

days total20

Semi-infinite, 

continuous 

Gases Produced H2, O2 H2, O2 H2, CO2 H2, CO2 H2, CO2

 

17 “...the immobilization technique allowed us to increase the cell density about 130-fold compared to a 

suspension culture” Kosourov, Sergey N. and M. Seibert. Hydrogen Photoproduction by Nutrient-Deprived 

Chlamydomonas reinhardtii Cells Immobilized within Thin Alginate Films under Aerobic and Anaerobic 

Conditions. NREL. Colorado. 16 June 2008. p 8. (NREL)

18 Calculated based on the molar ratios of photosynthesis and the rate of photosynthesis required to support 3% 

light conversion into biomass.

19 Operation in aerobic conditions is a possibility but is not postulated for this analysis. “Since this was a 

critical observation, we further investigated whether alginate-entrapped cells could produce H2 in the presence 

of air.... We in fact observed H2 gas production throughout most of the experiment. However, the final H2

yields were lower than the yields obtained under an argon atmosphere.” Kosourov, Sergey N. and M. Seibert. 

Hydrogen Photoproduction by Nutrient-Deprived Chlamydomonas reinhardtii Cells Immobilized within Thin 

Alginate Films under Aerobic and Anaerobic Conditions. NREL, Colorado, 16 June 2008. p 10.

20 Alternating cycles of 2-days growth/2-days H2 production with a total film lifetime of 180 days.
16 

B-1 B-2 B-3 B-4 B-5 

Theoretical & Assumed H2 Production Parameters

Assumed PAR21 44% 44% 44% 44% 71% 

Photons/H2 mol 4 4 4 4 11-15 

Theoretical Product 

Ratio

2 mol H2

22

1 mol O2

2 mol H2 

1 mol CO2

2 mol H2 

1 mol CO2

Experimental 

Product Gas Ratio

H2 production not yet demonstrated 100% H2

23 0.95 mol H2

0.05 mol CO2

Assumed Product 

Gas Ratio

2 mol H2

1 mol O2 

2 mol H2

0.86 mol CO2

2 mol H2

0.6 mol CO2

2 mol H2

0.05 mol CO2

STH24

(Max Theoretical)

Efficiency 12.2% 12.2% 12.2% 3% 6.5% 

Solar energy for 

Coincident Cell 

Growth25

3% 3% N/A N/A 3% 

STH Efficiency 

(Assumed Upper 

Bound)

9.2% 9.2% 5.2% (average 

over growth & 

production)

2.25% (average 

over growth & 

production)

3.5% 

STH Efficiency (Near 

Term Estimate)

2% 2% 1.3% 1.5% 1.5% 

Experimental STH 

Efficiency 

H2 production not yet demonstrated H2 not 

demonstrated

0.8% average 2.5%26

Bed Depth (cm) 10 10 10 1027 10 

Reactor Parameters 

H2 Rate (kgH2/day)28 1,111 1,111 1,176 1,176 1,111 

Number of 

Raceways

20 20 38 90 Production

2 Growth

54 

Raceway LxW: 

Raceway D

1090’x40’ 

x0.33’ 

1090’x40’ 

x0.33’ 

1090’x40’ 

x0.33’ 

1060’x40’ 

x0.33’ 

1090’x40’ 

x 0.33’

Reactor Area (m2) 80,968 80,968 151,754 352,070 216,228 

 

21 PAR: photosynthetically active radiation

22 Email from Pin-Ching Maness. (NREL) 10 October 2008. “...the gas output should be 1 mol O2 and 2 mol 

H2, with no CO2 at all.”

23 While CO2 must be produced by the reaction, it is dissolved in the aqueous media. Consequently no CO2 is 

experimentally observed in the headspace gas. However, in a potential future system, excess CO2 is expected 

to exceed the saturation limit and collect as a gas in the reactor bed headspace.

24 STH Efficiency =Solar to Hydrogen conversion efficiency = ratio of hydrogen net energy produced (lower 

heating value) to total solar energy incident on reactor bed.

25 This only applies to Chemostat II systems (section 4.1.4), where there is no separate growth stage.

26 Akkerman, Ida, Marcel Janssen, Jorge Rocha, Rene H. Wijffels. Photobiological hydrogen production: 

photochemical efficiency and bioreactor design. Intl. Journal of Hydrogen Energy. Vol. 27, p 1195-1208, 

2002.

27 Thinness of film substrate makes this depth unnecessary for photon capture, but it is still necessary for 

system temperature modulation.

28 Beds are sized to meet 1000 kgH2/day out of PSA. PSA recovery is determined by input hydrogen + 

contaminant gases for all beds, some of which are in growth phase. Calculated by PSA spreadsheet model and 

confirmed by gas separation experts at UOP LLC. H2 recovery for B1, B2, and B5 estimated at 90%. H2

recovery for B3 and B4 estimated at 85%.
17 

2.1 Nutrients

A wild algal colony reaches a final concentration in the photoautotrophic growth stage of >1 

g/L if allowed to grow to stationary phase. Thus, each system must be diluted through algae 

filtering to achieve a cell density of 0.2 g/L. The bed depth is chosen so that, following 

transition to the hydrogen production mode, the absorbance achieves the highest yield of 

hydrogen at this density. For the chemostat systems, a final cell concentration of 0.2g dry 

weight/L (WT cells) was recommended by NREL before transition to hydrogen production 

mode29. This concentration can be achieved either by dilution of the stationary phase culture 

or by earlier commencement of transition to hydrogen producing conditions. 

The concentration of algae is important because it determines how several other items are 

sized in our analysis. It impacts the amount of nutrients and input gases required. Later in 

the engineering parameters section we will analyze how it can affect pump and valve sizes 

as well.

2.1.1 Chlamydomonas reinhardtii Nutrients

In most laboratory environments, Chlamydomonas nutrient requirements are fulfilled by the 

tris-acetate-phosphate (TAP) medium with Hunter Trace Elements or other custom mixes of 

common lab chemicals. While ideal for a laboratory where these chemicals are readily 

available and only used in small quantities, on a production scale the costs are 

extraordinarily expensive. Even in large quantities and at technical grade, the 

manufacturing that goes into laboratory quality chemicals substantially increases the price. 

A cost breakdown of the TAP medium can be seen in Figure 2-2. Prices are based on 

quotes from Spectrum Chemical Corporation and Eastern Chemical Corporation. 

 

29 What we desire is an appropriate matching of organism density and bed depth so that full photon capture is 

achieved. Full photon capture will lead to maximum hydrogen production. Amos suggests that 0.2g/L is a 

feasible and practical organism density. Consequently, we have adopted that density and have adjusted bed 

depth based on the specific organism absorption properties to achieve full photon capture.
18 

Figure 2-2. TAP Medium Costs Compound Symbol Price Quantity Concentration Price Cost/ Liter

Tris (HOCH2)3CNH2 $13,480.00 907 kg 2.43 g/L $14.86/kg 0.03612

Ammonium Chloride NH4Cl $102.00 23 kg .4 g/L $4.5/kg 0.00180

Magnesium Sulfate, Heptahydrate MgSO4.7H2O $2,160.00 500 kg .1 g/L $4.32/kg 0.00043

Calcium Chloride, Dihydrate CaCl2.2H2O $583.00 50 kg .05 g/L $11.66/kg 0.00058

Potassium Phosphate, Dibasic K2HPO4 $216.00 23 kg .108 g/L $9.52/kg 0.00103

Potassium Phosphate Monobasic KH2PO4 $3,300.00 225 kg .056 g/L $14.67/kg 0.00082

Acetic Acid CH3COOH $558.00 208 L 1. ml/L $2.68/L 0.00268

EDTA Disodium Salt Na2EDTA $1,402.74 12 kg 50. g/L $116.9/kg 0.00584

Zinc Sulfate, Heptahydrate ZnSO4.7H2O $709.80 12 kg 22. g/L $59.15/kg 0.00130

Boric Acid H3BO3 $531.54 12 kg 11.4 g/L $44.3/kg 0.00050

Manganese Chloride, 4- Hydrate MnCl2.4H2O $1,227.78 12 kg 5.06 g/L $102.32/kg 0.00052

Ferrous Sulfate, Heptahydrate FeSO4.7H2O $865.44 12 kg 4.99 g/L $72.12/kg 0.00036

Cobalt Chloride, Hexahydrate CoCl2.6H2O $2,683.50 12 kg 1.61 g/L $223.63/kg 0.00036

Cupric Sulfate, Pentahydrate CuSO4.5H2O $716.82 12 kg 1.57 g/L $59.74/kg 0.00009

Ammonium Molybdate, Tetrahydrate (NH4)6Mo7O24.4H2O $2,509.13 12 kg 1.1 g/L $209.09/kg 0.00023

Total Cost TAP Medium/ Liter 0.0435

Total Cost Hunter Elements/Liter 0.0092

Total Cost/ Liter Solution 0.05267

Total Cost/ Reactor (B1) $427,037.50

Total Cost/ Year

(assuming 12 reactor changes per yr) $5,124,450.00

TAP Medium

Hunter Trace Elements 

Due to the prohibitive cost of the TAP medium, we investigated alternative nutrient sources. 

Fish aquaculture provides an excellent case study for large-scale algae production since 

algae is the primary food source for most fish, and an effective nutrient source is integral to 

the success of the culture. The standard source of nutrients for algal based aquaculture is 

liquid or solid farm fertilizer. The most effective fertilizers are ones that contain all three of 

the major fertilizer nutrients: Potassium, Phosphorous, and Nitrogen. Commercial fertilizers 

also come with a wide variety of available trace elements that can be added in any number 

of custom blends. The prices for commercial fertilizers that would be adequate for effective 

algae growth are substantially cheaper than those of the TAP medium or other similar 

media. Information from the US Dept. of Agriculture (USDA) indicates the cost of 

fertilizers ranges generally between $150 and $550 per ton, depending on the type of 

fertilizer and the quantity being purchased. Some of the more common fertilizers and their 

prices over the past seven years can be seen Figure 2-3. 

Figure 2-3. Fertilizer prices per ton since 2000.

 Year Month

Anhydrous 

ammonia

Nitrogen 

solutions (30%)

Urea 45-46% 

nitrogen

Ammonium 

nitrate

Sulfate of 

ammonium

Super-phosphate 

44-46% 

phosphate

Diammonium 

phosphate 

(18-46-0)

Potassium 

chloride 60% 

potassium

2000 Apr. $227 $131 $200 $194 $167 $233 $240 $165

2001 Apr. $399 $189 $280 $260 $192 $236 $244 $170

2002 Apr. $250 $127 $191 $195 $187 $221 $227 $164

2003 Apr. $373 $161 $261 $243 $195 $243 $250 $165

2004 Apr. $379 $178 $276 $263 $205 $266 $276 $181

2005 Apr $416 $215 $332 $292 $244 $299 $303 $245

2006 Apr $521 $232 $362 $366 $266 $324 $337 $273

2007 Apr $523 $277 $453 $382 $288 $418 $442 $280

According to US Patent #5,567,221 from OMS Investments Inc., the suggested aquaculture 

algae nutrient rates for the three major fertilizer nutrients is 100 pounds/acre per production 

cycle for nitrogen, 800 pounds/acre per production cycle for phosphorous, and 50 

pounds/acre per production cycle for potassium. Average production cycles vary depending 

on the species of fish from as much as 18 months to as little as 6 months. If we average that 
19 

range to a year, and estimate the algal concentration for these fertilizer quantities we can 

compute the necessary fertilizer for our application. 

The biodiesels industry indicates concentrations of 1g/L to 10 g/L of algae in solution are 

normally achieved. Taking the average we assume that the nutrient rates suggested by the 

OMS patent are based on an algae concentration of 5g/L, compared to our colony 

concentration of 0.2g/L. The bed depth in the OMS patent is 1m compared to our pond 

depth of 0.1m. The rate of nutrients of interest is g Nutrient/g algae/m2/day. In order to 

compute this we first convert the rate given to g/m2/day using the conversion factors below. 

Conversion Factors

4046.9 m2/acre

365 days/yr

453.592 g/lb

 Nitrogen Potassium Phosphorus

lb/acre/yr 100 50 800

g/m2/day 0.031 0.015 0.246

Then using the concentration values and bed depths we can compute the g algae/m2 for the 

reference and for our plant. 

Conversion Factors 

Volume 1000 L/m3

Bed Depth (Aqua) 1 m 

Bed Depth (Photobio) 0.10 m 

 Aquaculture Photobio

g/L 5 0.2

g/m2 5000 20

Lastly, we compute the daily mass ratio of nutrients to algae in the aquaculture then convert 

to an areal density given the Photobiological algae concentration. This leads us to the 

values provided in the Biological Parameters table above (Figure 2-1) for colony growth. 

Naturally, this amount would be less at the start of the colony. However, given the 

extremely small level of nutrients and the short growth time, we took this nutrient rate as a 

constant value during the growth phase. During H2 production the nutrient rates decrease as 

specified by other sources.

Figure 2-4. Nutrient Rates Computed

 Nutrient --> Nitrogen Potassium Phosphorus

Aquaculture g Nutrient/m2/day 0.031 0.015 0.246

Aquaculture/Photobio g Nutrient/g algae/day 6.142E-06 3.071E-06 4.913E-05

Photobio g Nutrient/m2/day 1.228E-04 6.142E-05 9.827E-04
20 

If we average the production cycle range to a year with rates computed, the cost of nutrients 

drops significantly when compared with the TAP medium. An example of possible nutrient 

cost for the B-1 system using three different fertilizers, each rich in one of the nutrients, is 

shown in Figure 2-5. The costs were calculated using the Nitrogen solutions, Super 

Phosphate, and Potassium Chloride prices from the USDA. While costs may vary and more 

accurate testing of the adequate fertilizer mix and the appropriate rate still need to be done, 

these costs are indicative of the cost for using fertilizer as a nutrient source. 

Figure 2-5. Sample Fertilizer Costs

Nutrient Cost/Lb Lb/Acre/Yr. Cost/Acre/Yr. Cost/1 TPD 

Module/Year

Nitrogen $0.14 0.4 $0.06 $0.80 

Potassium $0.14 0.2 $0.03 $0.40 

Phosphorous $0.21 3.2 $0.67 $9.63 

Total $0.49 3.8 $0.76 $14.65 

It should be noted that this price and nutrient breakdown are merely representative. In 

hydrogen production mode, a fertilizer without phosphorous would be used for the B-4 

system, while in the B-5 system, a fertilizer that is nitrogen deprived would be ideal. 

Additionally, small amounts of sulfur would be added to the B-4 system in order to maintain 

the cell during long hydrogen production periods. Since the cost of nutrients is 

insignificant, more accurate costing of each nutrient system was not necessary for this 

report. 

2.1.2 Bacteria Nutrients

Cyanobacteria (blue-green algae) are a diverse branch of prokaryotic oxygenic 

photoautotrophs that, like algae, oxidize water as electron donor and fix CO2 using light. 

They produce energy during the night by consumption of energy reserves, mainly glycogen, 

under aerobic conditions (respiration) or anaerobically (fermentation). They differ from 

algae in using a phycobilin pigment system for antenna. There are many thousands of 

different genera and strains with vastly different requirements for water, salinity, pH, 

temperature and nutrients. The native cyanobacterial strain serving as background for the 

genetic alterations considered above in B-2 is Synechocystis PCC6803. This unicellular 

strain grows in freshwater media similar to that used with C. reinhardtii. In our analysis, 

Cyanobacteria nutrients and nutrient rates are the same as for the C. reinhardtii. 

2.1.3 PNS Bacteria Nutrients

Purple Non-Sulfur (PNS) photosynthetic bacteria are very different organisms than algae 

and cyanobacteria and have different nutrient requirements. In previous studies concerning 

PNS, the primary source of nutrition was a combination of laboratory chemicals; however, 

organic acids such as acetic acid combined with agricultural fertilizer containing the three 

major macronutrients are hypothesized for mass production. Acetic acid consumption 
21 

during hydrogen production of PNS bacteria has been demonstrated experimentally30. 

Based on the assumed PNS hydrogen production reaction31

Figure 2-6

of C2H4O2 + 2H2O + Light 

Energy => 2 CO2 + 4 H2, 4 moles of hydrogen gas are formed for each mole of acetic acid. 

Using this ratio, along with the ratio of acetic acid to PNS cell mass (modeled as C5H8O2) 

needed for the assumed 3% of photons dedicated to growth, we can calculate the overall

acetic acid consumption rate shown in . 

Figure 2-6. Acetic acid Consumption by PNS

Molar Ratio of Acetic Acid to PNS Mass (Section 3.5) 3

Molecular Weight of Acetic Acid g/mol 60.032

Molecular Weight of PNS Mass g/mol 100.064

Mass of PNS Grown (3% of photons) kg/day 468

Acetic Acid Needed for Growth kg/day 842

Molar Ratio of Acetic Acid to H2 0.25

Molecular Weight of H2 g/mol 2.016

H2 Produced kg/day 1111

Acetic Acid Needed for H2 Production kg/day 8272

Total Acetic Acid Needed kg/day 9113

PNS Acetic Acid Consumption (Based on Theoretical Mass Balance)

According to data from the Chemical Journal (Korea) the international market price for 

acetic acid is $0.595/kg32. As an alternative, in an integrated system, the waste product of a 

fermentor can potentially be used as an organic acid feedstock for the PNS bacteria.

2.1.4 pH of Solution 

The pH of the solution in the reactor beds should be maintained near 7 to ensure optimal 

organism growth. In systems B-3, B-4 and B-5, there will be trapped CO2 released by the 

organisms in the process water. In the B-3 and B-4 systems this CO2 is needed as a 

feedstock for the algae during the growth stage to return the cell concentration to a level 

high enough to support the respiration needed to keep the system anaerobic during hydrogen 

production. Approaches to maintaining targeted pH levels have not been considered in 

detail within this report. Further investigation of the effects of pH should be considered in 

future work. 

2.1.5 Carbon Dioxide 

In addition to fertilizer and acetate, these organisms utilize carbon dioxide to make both 

structural biomass and energy reserves. The gas rate requirements have been determined 

using the stoichiometric ratios of photosynthesis. Additional CO2 will only need to be 

added to the B-1 and B-2 systems to support the minimal growth required to keep the 

system alive. Systems B-3 and B-4 will be producing all of the CO2 necessary for the 

 

30 Barbosa, Maria, Johannes Tramper Jorge Rocha, Rene H. Wijffels. Acetate as a carbon source for hydrogen 

production by photosynthetic bacteria. Journal of Biotechnology. Vol 85, p 25-33. 2000.

31 Akkerman, Ida, Marcel Janssen, Jorge Rocha, Rene H. Wijffels. Photobiological hydrogen production: 

photochemical efficiency and bioreactor design. International Journal of Hydrogen Energy. Vol 27, p 1195-

1208. 2002.

32 Lee, Sang Y. Plastic Bacteria? Progress and prospects for polyhydroxyalkanoate production in bacteria. 

Trends in Biotechnology Vol 14, p 431-438. 1996.
22 

system, and the B-5 system will utilize acetate as its carbon source. The amount of carbon 

dioxide required was checked against the amount soluble in water. The amount of CO2 used 

should be well below saturation to avoid excess CO2 bubbling out into the H2 product gas. 

Using the CRC Handbook of Chemistry and Physics, we obtained the solubility of CO2 in 

water. We will maintain a CO2 partial pressure that will assure adequate CO2 supply. In the 

B-3 and B-4 systems the CO2 produced by the algae will be far above the solubility 

threshold, and the gaseous CO2 produced will be filtered out through PSA and added to the 

system after the dissolved level of CO2 drops through algae consumption. Since each 

system consists of numerous beds, their growth and production cycles will be appropriately 

staggered to ensure that there are always beds in need of the CO2 that is being filtered by the 

PSA. A constant slip stream of water will be removed and run through a CO2 saturation 

vessel and returned to raceways that are low on CO2. The only addition of CO2 necessary 

for B-3 and B-4 is the initial amount needed at the initial colony growth. B-5 has no CO2

needs. 

The B-1 and B-2 organism needs for normal growth conditions are 31.66mg/L/day and for 

reduced growth conditions are 146.57 mg/L/day18. Figure 2-7Figure 2-7. Carbon Dioxide 

Solubility in Water shows that these CO2 concentrations are well below the maximum 

solubility of CO2 in water at one atmosphere and that the CO2 input will remain in solution. 

The CO2 needs for sustainment for the entire system amounts to only 257kgCO2/day 

(17,560ft3/day at 1atm). The CO2 can be readily provided by maintaining a CO2 partial 

pressure of 0.5psia. This partial pressure will also minimize the CO2 dilution of the product 

H2 in the headspace of the reactor. 
23 

Figure 2-7. Carbon Dioxide Solubility in Water 

0

500

1000

1500

2000

2500

3000

3500

0 10 20 30 40 50

mg CO2/ liter water

Temperature oC

CO2 Solubility in Water 

at 1 bar (0.99 atm)

CO2 max solubility vs temperature

146.57 mg CO2 / L / day

31.66 mg CO2 / L / day

2.2 O2-tolerant Hydrogenase (B-1 & B-2)

The B-1 and B-2 systems utilize organism mutants that are O2-tolerant, with B-1 using the 

algae strain C. reinhardtii and the B-2 system using the cyanobacterium Synechocystis

PCC6803. The O2-tolerant hydrogenase pathway in both organisms is comprised of 

theoretical mutants with truncated Ch1 antennae (Light Harvesting Complex, or LHC, 

mutants) to minimize excess absorption of solar energy beyond what is needed for maximal 

H2 production. Two research groups have reported mutants in C. reinhardtii that have 

reduced Chl content associated with truncated antenna complexes.33,

34 In both cases, the 

cells exhibit a reduced optical cross section and a higher growth rate at full solar light 

intensity. Factors of two have been reported. However, in both cases, reports on hydrogen 

production from these mutants have not been published, since research is ongoing. 

A number of oxygen tolerant hydrogenases have been identified in bacteria as potential 

enzymes in algal hosts; bioengineering has yet to produce a mutant that exhibits stable H2

output over extended periods. 

 

33 Melis, A. (2007) Photosynthetic H-2 metabolism in Chlamydomonas reinhardtii (unicellular green algae), 

Planta 226, 1075-1086.

34 Mussgnug, J. H., Thomas-Hall, S., Rupprecht, J., Foo, A., Klassen, V., McDowall, A., Schenk, P. M., 

Kruse, O., and Hankamer, B. (2007) Engineering photosynthetic light capture: impacts on improved solar 

energy to biomass conversion, Plant Biotechnology Journal 5, 802-814.
24 

The current best solar-to-hydrogen (STH) energy conversion efficiency of non-tolerant C. 

reinhardtii is ca. 0.8% for immobilized cells35. The theoretical prediction for algae 

assuming 100% flux of photoelectrons going solely into H2 production and without any loss 

or competition with other processes predicts an STH of ~12.2%. However, a certain amount 

of energy must be diverted for homeostasis, the regulation of the general health and 

functioning of the cell, regardless of the outside environment. Allowing for energy losses 

required for homeostasis, but assuming no competition with other pathways for utilization 

of light energy, our analysis assumes that hypothetical mutant B-1 algae will produce 

hydrogen at 9.2% STH efficiency, while growing at approximately 1⁄4 of the normal growthonly doubling rate (i.e. the doubling rate of algae that is not producing H2). Because the 

growth rate during production is assumed to be 1⁄4, the carbon dioxide and nutrients during 

this condition is also assumed to be 1⁄4. 

For the B-2 system, Cyanobacteria do not produce significant levels of hydrogen upon 

illumination, even those strains which possess a native hydrogenase. However, 

hydrogenase-containing strains can produce hydrogen upon incubation under dark anaerobic 

conditions which induces autofermentation. In order to produce significant levels of 

photohydrogen with cyanobacteria, an engineered strain will need to be derived in which the 

native [NiFe]-hydrogenase is replaced with one capable of utilizing the same photoreductant 

produced in algae, ferredoxin. Though deemed possible, no such bacteria have been 

reported to date in the literature. Therefore, extrapolations made herein are based on 

predictions provided by the NREL staff. The algae assumptions of a theoretically possible 

12.2% STH for the B-1 pathway are also applied to the B-2 pathway. 

Pathways B-1 and B-2 are expected to produce hydrogen indefinitely unless the colony is 

spoiled by external factors. Since these mutants are still in a research stage, our analysis 

will assume that the amount of energy spent on hydrogen production can be regulated 

somewhat by controlling CO2 or ferredoxins (controlling how much energy goes to cell 

maintenance and cell growth).36

Figure 2-1

 Although the organism is different for photobiological 

system B-1 and B-2, we assume their growth rates and calculated parameters to be identical 

in , as the systems are future projections of yet undeveloped strains. 

Cyanobacteria may have an advantage to algae as their native hydrogenase is less sensitive 

to oxygen7. However, as noted above they produce H2 most efficiently in the dark and 

would require an algal-like hydrogenase if photohydrogen was the primary goal. As more 

data become available for these systems, the STH performance can be modified to assess the 

impact on hydrogen production capabilities and system costs. 

 

Photobiological Systems B-1 and B-2 are modeled by the following reactions: 

H2 Production: 2 H2O + Light Energy  2H2 + O2

Respiration: 6O2 + C6H12O6  6CO2 + 6H2O 

Photosynthesis/Growth: 6CO2 + 6H2O + Light Energy  C6H12O6 + 6O2

The inputs and outputs of the system are depicted in Figure 2-8. 

 

35 Kosourov, Sergey N. and Michael Seibert. Hydrogen Photoproduction by Nutrient-Deprived 

Chlamydomonas reinhardtii Cell Immobilized within Thin Alginate Films under Aerobic and Anaerobic 

Conditions. Basic Sciences Center, NREL, Golden Colorado. 16 June 2008.

36 Telecom. 09JUL08 Mtg Minutes. Technoeconomic Analysis of Bio-Hydrogen Program. NREL Contract # 

AFH-8-88601-01. Revision 10 July 2008.
25 

Figure 2-8. Oxygen-Tolerant Hydrogenase Process

GASES

LIQUIDS

SOLIDS

LEGEND

O2-tolerant

Hydrogenase

B-1, B-2

Inputs Outputs

66% H2 by mol, 

34% O2 by mol

Water+Algae Waste Algae 

Water

Water

Nutrients

CO2

The product gases from this system are hydrogen and oxygen. In certain ratios, those gases 

can be a highly combustive mixture. The gas mixture would be located above the reactor 

bed. The engineering analysis portion of this report addresses ways to separate this mixture 

to obtain the pure product hydrogen. 

2.3 Sulfate-permease Green Algae (B-3)

Like the B-1 pathway the sulfate-permease pathway utilizes algae mutants with truncated 

antennae. However, the algae in this pathway lack the mutated O2-tolerant hydrogenase of 

B-1. Hydrogen production is controlled through genetic mutation of the chloroplast sulfur 

uptake mechanism, reducing the rate of oxygenic photosynthesis. Sulfur is necessary for the 

regeneration of proteins damaged by photo-oxidation. Over time, the build-up of damaged 

photosynthetic proteins in the chloroplast will lead to a reduction in oxygen evolution. As 

the rate of oxygen evolution falls below the rate of oxidative respiration, the system will 

become anaerobic, leading to the induction of the hydrogenase pathway and continuous 

hydrogen production. The level of respiration needed to eliminate the high levels of oxygen 

produced by the water splitting hydrogen production reaction will necessitate a large amount 

of starch that must be produced by the cells prior to hydrogen production. Therefore, this 

organism is operated in a cyclic manner, with: (1) a growth/regeneration leading to starch

buildup and (2) a hydrogen production phase which is anaerobic. In order to achieve the 

required level of starch mass, the starting concentration of the cells at the end of the 

hydrogen production phase will be 0.85g/l and will increase rapidly over the course of the 

regeneration phase to ensure enough mass exists to keep the system anaerobic when 

subsequently producing the necessary volume of hydrogen gas. Over time, however, the 

lack of starch production and the loss of cell mass due to respiration keeping the system 
26 

anaerobic will inhibit cellular function, necessitating the regeneration phase where the 

system is bubbled with CO2 and air. For phase transition, adding small amounts of air to the 

system will shut down the hydrogenase enzyme, ceasing H2 production and directing 

photosynthesis towards algae mass production. 

During the production phase, the carbon dioxide production from respiration will rapidly 

saturate the process water and be released into the headspace. This CO2 will be removed 

from of the product gasses using a PSA and then returned to other beds in the system that 

are in growth phase, and thus need CO2 (the individual raceways will be staggered so that 

not all of them will be in the same phase of the cycle at once). Data has indicated that a 

system like this can engage in continuous hydrogen production for 80-100 hours. 

Experimental data are lacking for hydrogen production rates since research is ongoing and 

data has yet to be published. Theoretically, while in hydrogen production mode, the sulfate 

permease mutant could convert 100% of photoelectrons into H2 production without any loss 

or competition with other processes. However, the maximal 12.2% STH efficiency of the 

photohydrogen production must be averaged over both phases, and thus is reduced by a 

factor of 3/7, given the need for a 3-day hydrogen production period followed by a 4-day 

recovery and regeneration period. 

Photobiological System B-3 is modeled by the following reactions: 

H2 Production Phase 

H2 Production Reaction: 2 H2O + Light Energy  2H2 + O2

Respiration Reaction: 6O2 + C6H12O6  6CO2 + 6H2O 

Growth/Regeneration Phase:

Photosynthesis Reaction: 6CO2 + 6H2O + Light Energy  C6H12O6 + 6O2

Respiration Reaction: 6O2 + C6H12O6  6CO2 + 6H2O 

The inputs and outputs of the system for both production and growth phase are depicted in 

Figure 2-9. 
27 

Figure 2-9. Sulfate Permease Process

GASES

LIQUIDS

SOLIDS

LEGEND

Sulfate Permease

B-3

Inputs Outputs

Water H2 and CO2

Sulfate Permease

B-3

Inputs Outputs

O2 

Water

Nutrients

CO2

Hydrogen Production Phase

Growth Phase

Water

Air

2.4 Immobilized sulfate-deprived green algae (B-4)

In this photobiological pathway the basic system concept utilizes mutant, truncated-antenna 

Chlamydomonas reinhardtii algae. It involves use of a designated bed for algal growth 

followed by immobilization of the algae by attachment to a thin film in the growth bed. The 

film with immobilized algae is transferred to a production bed for the H2 production mode, 

which is initiated by sulfur deprivation. Sulfur deprivation leads to the gradual and 

reversible reduction of O2 evolution in algae. As the rate of O2 evolution falls below the rate 

of oxidative respiration, the system goes anaerobic, leading to H2 production. The benefit of 

thin film immobilization is a substantial increase in duration of the H2 production mode as 

compared to non-immobilized sulfate-deprived algae systems due to the reduced energy 

requirements of the cells for movement and increased cell density. Unlike the B-1 and B-2 

pathways, the immobilized system has been successfully tested in a laboratory environment. 
28 

Data for the immobilized, sulfate-deprived green algae are based on an NREL study by 

Kosourov and Seibert37. The NREL study utilized a thin alginate film as the immobilization 

substrate; however other types of films have also been used in analogous applications. For 

example, nitrogen-fixing blue-green algae have been attached using their heterocyst cells to 

cellulose sheets and to spunbond and meltblown polypropylene web sheets for use in 

stimulating agricultural plant growth38. In order to improve the affinity of the algae to the 

polypropylene film, the film surface energy will probably need to be increased to a level of 

40 dynes/cm using corona treatment, or an alternative material having higher inherent 

surface energy, such as polyester can be used. With any man-made polymer, it is ultraimportant to steam clean or otherwise treat the material to eliminate the manufacturing 

process substances that would be deleterious to the algae. 

In other research work, cyanobacteria have been immobilized on polyurethane foam sheets 

(as well as alginate) for hydrogen production purposes at King's College in London39. As 

noted therein, polyurethane foam was also used to immobilize the algae Chlorella vulgaris. 

In this report, we have chosen to use meltblown polypropylene web sheets for minimum

cost immobilization. The Kosourov report uses an alginate/plastic screen configuration. 

This alginate has advantages if the immobilized algae are later used in a fermentation 

process, since the alginate film will also ferment. However, the higher cost of the screen 

along with the cost associated with impregnating the screen with alginate increases the 

levelized hydrogen cost of the B-4 system. Conversations with NREL have also indicated 

that on substrates other than alginate, hydrogen production periods of up to six months can 

be maintained with the addition of minimal amounts of sulfur40. This increase in hydrogen 

production time will reduce the labor requirement for film replacement. 

As with the B-3 system, there will be a hydrogen production phase followed by a 

growth/regeneration phase in which the concentration and mass of algae will increase 

enough to allow adequate starch for the respiration necessary to keep the system anaerobic 

during the production phase. This growth/regeneration phase will be triggered by the 

addition of sulfur to the nutrient medium. Sulfur deprivation will then be used to trigger the 

hydrogen production phase. At the end of a two day hydrogen production phase, the 

concentration of algae will be at 1.81/L, the concentration needed for the necessary mass 

growth to allow enough respiration for anaerobic hydrogen production. Also like the B-3 

system, the beds will be staggered to allow excess CO2 separated through PSA to be added 

to a slip stream of water removed from raceways during the regeneration phase. After 180 

days of cycling between growth and production, the algae films will need replacement from 

mats being created in the two designated growth beds. 

 

37 Kosourov, Sergey N. and Michael Seibert. Hydrogen Photoproduction by Nutrient-Deprived 

Chlamydomonas reinhardtii Cell Immobilized within Thin Alginate Films under Aerobic and Anaerobic 

Conditions. Basic Sciences Center, NREL, Golden Colorado. 16 June 2008.

38 U.S. Patent 4,879,232, Nov 7, 1989, "Multilayered Structure Containing Immobilized Blue-Green Algae”.

39 Muallem, Bruce, and Hall, "Photoproduction of H2 and NADPH2 by Polyurethane Immobilized 

Cyanobacteria", King's College, London.

40 Private communication with Dr. Mike Seibert of NREL, 29 May 2009.
29 

Theoretically, the algae can be separated from the binder and rejuvenated but, since the 

algae growth costs are low, it is not worth the complexity of the separation system. Figure 

2-10 shows the growth and production modes of this pathway. 

Figure 2-10. Immobilized Sulfur Deprived Process

GASES

LIQUIDS

SOLIDS

LEGEND

Sulfur Deprived 

Immobilized

B-4

Inputs Outputs

Water H2 and CO2

Sulfur Replete 

Immobilized

B-4

Inputs Outputs

O2 

Water

Nutrients

CO2

Hydrogen Production Phase

Growth Phase

Water

Immobilized Film 

Creation

B-4

Inputs Outputs

Nutrients O2 

CO2

Water

Film Algae Film

Algae/Film Growth Bed Hydrogen Production Bed

In this system we assume the same growth rate as other Chlamydomonas systems. To 

ensure that the growth beds are fully utilized, we will use two algae film/growth beds to 

produce algae for the 68 production beds. The H2 production cycles of each bed will be 

staggered appropriately, so that each day the growth beds will be utilized to replenish a 

spent production bed. 

Compared to O2-tolerant algal systems (B-1, B-2), cellular maintenance energy 

requirements are substantially lower due to the immobilization of the algae on thin films. 

The theoretical STH efficiency is 3%; however, the laboratory tests have only achieved

0.8% STH efficiency37. Since this is an actual data point, it is not reasonable to compare it 

to other systems that only have theoretical values. For comparison purposes we will use the 

STH efficiency of the assumed reactor operation of 2.25%, due to the one day of production 

for every three days of growth. 
30 

Photobiological System B-4 is modeled by the following reactions: 

H2 Production Phase 

H2 Production Reaction: 2 H2O + Light Energy  2H2 + O2

Respiration Reaction: 6O2 + C6H12O6  6CO2 + 6H2O 

Growth Phase: 

Photosynthesis Reaction: 6CO2 + 6H2O + Light Energy  C6H12O6 + 6O2

Respiration Reaction: 6O2 + C6H12O6  6CO2 + 6H2O 

2.5 Purple non-sulfur bacteria (B-5)

Unlike the carbon-fixing organisms in the previous systems, these bacteria are 

photoheterotrophic, nitrogen-fixing organisms. Under anaerobic conditions and using 

simple organic acids or glycerol as an electron donor, these bacteria produce hydrogen gas 

from the light driven, ATP-dependent reduction of protons to H2 in the absence of 

nitrogen41. In this mode they require removal of N2 and O2 and require a supply of organic 

acids. The photosynthetically active radiation, or PAR, for purple bacteria is 400 – 950 nm, 

compared to 400 – 700 nm for green algae. However, the PNS organism requires 11 to 15 

photons (depending on wavelength) per hydrogen molecule rather than the 4 photons 

required in the algal systems. As a result of the wavelength energy span and the required 

photon ratio, the maximum solar-to-hydrogen conversion efficiency is 6.5%. Due to the 3% 

of photons needed for cell maintenance, the assumed reactor STH is 3.5%. 

Since the PAR wavelength of PNS is shifted relative to algae, there is an opportunity to 

capture a fuller spectrum of light energy by “stacking” an algae system on top of a PNS 

system. This concept will be further described in Part IV (System Integration). 

Photobiological System B-5 is modeled by the following reactions: 

H2 Production Reaction: C2H4O2 + 2 H2O + Light Energy  2CO2 + 4H2

Fermentation Reaction42

Growth Reaction

: 2C6H12O6 + 2H2O  5C2H4O2 + 2CO2 + 4H2

43: 3C2H4O2  (C5H8O2)n + CO2 + 2H2O 

The inputs and outputs of the system are depicted in Figure 2-11. 

 

41 Akkerman, Ida, Marcel Janssen, Jorge Rocha, Rene H. Wijffels. Photobiological hydrogen production: 

photochemical efficiency and bioreactor design. International Journal of Hydrogen Energy. Vol 27, p 1195-

1208. 2002.

42 The fermentation reaction produces additional hydrogen as well as organic acids (here modeled as acetate). 

However, it doesn’t produce these in appreciable quantities and is thus ignored in this analysis.

43 The actual PNS growth reaction is complex and involves ammonia (NH3) and multiple intermediaries. For 

simplicity and clarity, we base our analysis on the displayed simplified net reaction. This photosynthesic 

growth reaction produces a variety of organic products, including the glucose utilized in fermentation. Ashing 

analysis (from Van Gemerden H. 1968. Utilization of reducing power in growing cultures of Chromatium. 

Arch. Mikrobiol. 64:111) shows that the molecular ratio is represented as (C5H8O2)n. 
31 

Figure 2-11. Purple Non-Sulfur Bacteria Process

GASES

LIQUIDS

SOLIDS

LEGEND

PNS

B-5

Inputs Outputs

H2, CO2

Water+Bacteria

Waste Bacteria

Water

Water

Nutrients, Acetic Acid

3.

The production plant has been divided into the Photobiological Reactor Bed and 4 auxiliary 

or balance of plant (BOP) subassemblies, namely, 

Engineering Parameters 

• Organism Feed Subassembly, 

• Recycle Subassembly, 

• Gas Capture Subassembly, and 

• Control System Subassembly 

The interfaces between the primary system and its BOP subassemblies are shown in Figure 

3-1. Each subassembly is depicted in a different color. The Photo Bio Reactor Bed can take 

on several forms. The subassemblies have several components that will be identified in 

later sections of this document. 
32 

Figure 3-1. Simplified Photobiological Hydrogen Production Plant

Gas Capture

Subassembly

Photo Bio

Reactor Bed

Primary

System

Hydrogen

Pipeline

Power

Source

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Organism Feed

Subassembly

Recycle

Subassembly

Control System

Subassembly

PC

Instrumentation

A modular design philosophy was used in the design of the production plant. Each function 

that needs to be performed in the plant is treated as an independent subassembly and is sized 

appropriately so that multiple systems can be arranged in parallel within a production plant. 

The subassemblies are designed such that a single module produces 1,000kgH2/day at the 

boundary limits of the production facility. As will be discussed in the results section, the 

baseline cost estimates correspond to a 10 tonne/day plant consisting of ten 1,000kgH2/day 

modules. 

3.1 Photo Bio Reactor Bed (Primary System)

Several designs for reactors beds have been developed to enhance performance of these 

photobiological systems. Each system has both strengths and weaknesses. Our study 

reviewed four different forms for reactor beds. Each form was then applied to a specific 

organism to determine which was more or less suitable for the different algae and bacteria. 

The reactor bed form was selected based on the author’s initial judgment of cost 

effectiveness and because an analysis of all five pathways with all bed form options was 

beyond the scope of the project. 

3.1.1 Single Reactor Bed

The simplest form for a bed is the single reactor. In this bed there is one pond where all 

photobiological processes occur. The processes occur sequentially, not simultaneously. 

Thus, if algae are in the growth phase they cannot produce hydrogen and vice versa. The 

strengths and weaknesses of this system are listed in Figure 3-2. 
33 

Figure 3-2. Strengths and Weaknesses of Single Bed Reactors

Single Bed Reactor Design

Strengths Weaknesses

Organisms are working at their most 

efficient point during each process.

Batch process 

Single pond; less area, lower cost Time split between production and growth 

not most efficient.

Requires more area for a given hydrogen 

production

A conceptual design of the single reactor bed is shown in Figure 3-3. The blue box 

represents the reactor bed. The dotted line separates the different processes that occur in the 

same bed but at different times. The organisms transition from one photobiological process 

to another by a manipulation of the inputs and climate conditions. Sunlight is required 

during both photobiological phases. 

Figure 3-3. Conceptual Design: Single Bed Reactors

Sin g l e B ed R ea c t o r

Duration: Phase 2 = Phase 1

Growth

Phase 1 Phase 2

Production

Reactor Bed

3.1.2 Dual Reactor Bed

The next form for a bed is the dual reactor. In this reactor there are two ponds; one where 

all photobiological growth occurs and another where photobiological hydrogen production 

occurs. Because there are separate beds for each process, growth and production can occur 

simultaneously. The strengths and weaknesses of this system are listed in Figure 3–4. 
34 

Figure 3-4. Strengths and Weaknesses of Dual Bed Reactors

Dual Bed Reactor Design

Strengths Weaknesses

Simultaneous growth and production Cost of two reactor beds

Growth and production bed sizes can be 

optimized for maximum utilization

Batch process

No mixing of nutrients or gases from 

production and growth stages

Large flow pumps required under 

intermittent operation (transfer process)

A conceptual design of the dual reactor bed is shown in Figure 3-5. The blue boxes 

represent the reactor beds. The organisms are transferred from the growth reactor bed to the 

production bed when the colony reaches a given concentration. Although growth and 

production occur simultaneously this is still essentially a batch process. Sunlight is required 

during both photobiological phases. This dual bed concept is useful when growth and 

production conditions are different, such as when growth can occur aerobically but 

production cannot. 

Figure 3-5. Conceptual Design of Dual Bed Reactors 

Production

Reactor Bed

Dual Bed Reactor

Growth

Reactor Bed

Water 

+

Organism

Batch

3.1.3 Chemostat Reactor Bed

A chemostat system requires two reactor beds. Like the dual bed reactor, there is one pond 

where all photobiological growth occurs and another where photobiological hydrogen 

production occurs. Because there are separate beds for each process, growth and production 

occurs simultaneously. A chemostat system seeks to maintain a constant volume in the 

reactor bed. This is accomplished by a continuous slip stream of volume being removed 

while simultaneously being replenished from the growth reactor bed. The volume of slip 

stream is computed such that the hydrogen production is constant. Because of the constant 

volume criteria, chemostat reactors continuously grow algae and produce hydrogen (as long 

as sunlight is available). The strengths and weaknesses of this system are listed in 

Figure 3-6. 
35 

Figure 3-6. Strengths and Weaknesses of Chemostat Bed Reactors

Chemostat Bed Reactor Design

Strengths Weaknesses

Simultaneous and continuous growth and 

production

Large slip stream volumes of up to 25% /day

Isolation of growth and production products No means of selectively separating algae 

such that slip stream contains waste algae

Smaller pumps required and operate 

continuously (their preferred mode)

A conceptual design of the chemostat reactor bed is shown in Figure 3-7. Conceptually, the 

chemostat form looks identical to the dual reactor bed however the water + organism flow 

between the two reactors is continuous rather than intermittent. Sunlight is required during 

both photobiological phases. 

Figure 3-7. Conceptual Design of Chemostat Bed Reactors

Chemostat Bed Reactor

Production

Reactor Bed

Growth

Reactor Bed

Water 

+

Organism

Continuous

3.1.4 Chemostat II Reactor Bed

The Chemostat II’s chief characteristic is a single reactor bed that, after the initial growth of 

the colony, is used simultaneously for both growth and production. The system operates 

similarly to a chemostat system however it does so in a single reactor bed. During initial 

growth no production occurs. The bed is filled with the appropriate nutrients required for 

algae colony to achieve a target algal concentration. Afterwards, climate conditions are 

adjusted such that algae grow at a 1⁄4 of their normal rate and maintain cell functions while 

producing hydrogen. 

In H2 production mode, the organisms use a limited amount of their energy to complete cell 

maintenance functions and the remainder of the energy is spent producing hydrogen. By 

simultaneously doing growth and production, the Chemostat II system allows for continuous 

hydrogen production without having to stop and replenish algae. The strengths and 

weaknesses of this system are listed in 

Figure 3-8. 
36 

Figure 3-8. Strengths and Weakness of Chemostat II Bed Reactors

Chemostat II Bed Reactor Design

Strengths Weaknesses

Single pond; less area, lower cost Continuous supply of nutrient may result in 

diluted product stream

Lower nutrient requirement

A conceptual design of the Chemostat II reactor bed is shown in Figure 3-9. The Chemostat 

II looks similar to the single bed reactor, however growth and production occur 

simultaneously and continuously. Sunlight is required during both the growth and reduced 

growth with production processes. 

Figure 3-9. Conceptual Design of Chemostat II Bed Reactors

Reduced Growth

&

Production

Reactor Bed

C h emo st at II B ed R ea c t o r

Initial 

Growth

Duration: Phase 2 >> Phase 1

Phase 1 Phase 2

For this system to be feasible the number of days spent growing the culture must be 

insignificant when compared to the number of hydrogen production days. The colony 

concentration will increase slightly every hour and therefore a slip stream of the algal 

solution will be removed from the bed on a continuous basis. As that stream is being drawn, 

the reactor bed will be dosed with additional water and nutrients to maintain the colony at 

constant concentration, size, and pH levels. 

3.1.5 Reactor Bed Embodiment

The reactors forms were mapped against the biological parameters of each photobiological 

system and the best form was selected for each process. The forms selected are listed in

Figure 2-1. A complete bill of materials for each pathway is provided in section 4 of this 

report. Here is included a lengthy description of the physical embodiment of these reactor 

beds. The components of the reactor bed are described and materials selected for the bed 

are listed in this section. 
37 

A reactor bed could be open or closed to the environment. While an open pond would 

provide direct access to sunlight, it also leaves our algae colony susceptible to 

contamination and provides no means of collecting the production hydrogen. A closed pond 

with a translucent cover allows light to penetrate while protecting the algae from weather 

conditions (wind, temperature) that may harm them. It also contains any gases produced by 

the algae. For these reasons all reactor beds in this analysis will be of the closed pond 

variety. 

The reactor bed serves 4 basic functions: fluid containment, product gas containment, access 

to sunlight, and algal mixing (to avoid stratification). The bed can be a single monolithic 

structure, such as a pond, or several smaller beds, such as raceways, connected by a piping 

system. Since typical reactor bed areas are quite large (~20-87 acres for a 1,000kgH2/day), 

a modular concept was chosen where the reactor bed is comprised of smaller raceways. 

This facilitates ease of access, ease of maintenance, isolation of any contaminants, and 

economies of scale in equipment purchases. The reactor bed is sketched in Figure 3-10 and 

the components are listed in Figure 3-11. 

The reactor bed we have chosen is comprised of several raceways, each measuring between

1060’ and 1090’ in length, 40’ in width, and 10cm in depth. Raceways are designed in an 

oval shape so that a system of paddlewheels creates a flow in the bed to evenly disperse 

nutrients and prevent cell coagulation. Each raceway has inlet and outlet ports for the slurry 

and a separate outlet port for the product gases. Raceway widths are limited by the 

translucent film dimensions. The film is attached to a frame to maintain a gas-tight seal in 

the system. Driveways are interspersed between the raceways to provide vehicular access. 

The quantity of raceways is driven by daily hydrogen production rate and STH of the 

organism. For our systems there will be between 20 and 90 raceways. 

In concept B-4, immobilized organisms are grown on an inert mat which floats upon the 

surface of the raceway. Circulation of the water/nutrient mixture via paddlewheels is not 

needed nor is it feasible. Consequently, an oval reactor bed is not needed, and a “runway” 

configuration is instead used. A “runway” is merely a single side of the oval raceway, 

without the paddlewheels and without fluid connection between the two sides. 
38 

Figure 3-10. Reactor Bed Design

Inlet & Outlet Ports:

Makeup Water &

Nutrients

Gas Outlet Ports

Raceways

Pond Area

Algal Colony

Driveways

Transparent Film

Earthen Berms Raceways Pond Liner

Gas Tight Frame

Film Attachment Points Driveway

Shallow Earth

Trough Bed Depth

Paddlewheel

Paddlewheels

Inlet Ports: Makeup

Water & Nutrients

TOP VIEW

FRONT VIEW

Photo Bio

Reactor Bed

Primary

System

To Gas

Capture

To Recycle

From Algae

Feed

To Gas

Capture

To Recycle

From Algae

Feed

Figure 3-11. Reactor Bed Components 

Reactor Bed Components

Transparent Film

Pond Lining

Pond Edging

Paddlewheel & Motor

Gas Valves

Gas Ports (Flanges)

Water Valves

Water Ports (Flanges)

3.1.5.1 Films, Edging, and Lining

In the Front View one can see that the bed itself is not made with any solid materials but 

rather shaped from earthen berms and lined with a suitable pond liner material that will keep 

the system gas tight. Pond liner material is available in roll sizes of 50’ x 100’ on 

www.justliners.com. To achieve the necessary pond dimensions the material will be heatsealed to itself, which is a common practice when creating any kind of retention pond. 
39 

Above the raceways a low density polyethylene (LDPE) transparent film (6 mm) is used to 

allow sunlight in and keep the product hydrogen gas from escaping. A presentation from 

Daniel Blake at NREL concerning Hydrogen reactor development indicates that the 

hydrogen permeability coefficient for polyethylene is fairly low (near 200

cm3•mm/m2•atm•day)44. Based on our calculations, show in Figure 3-12, the volume of 

hydrogen lost through the film is negligible. 

Figure 3-12. Polyethylene Hydrogen Losses 

Units B1/B2 B3 B4 B5

Hydrogen Permeability Coefficient cm3

•mm/m2•

atm•day

200 200 200 200

Total Area m2 80,968 151,035 352,070 216,228

Film Thickness mm 6 6 6 6

Partial Pressure Differential atm 0.67 1.00 1.00 0.95

kg/cm3

 Hydrogen 0.00000009 0.00000009 0.00000009 0.00000009

H2 lost kg/day 0.16 0.45 1.06 0.61

% H2 lost from system % 0.016% 0.045% 0.106% 0.061%

Polyethylene Hydrogen Losses

According to conversations with Berry Plastics, the LDPE film allows 90% of available 

light through. It is available in rolls 56’ wide and more than 1000’ long and, therefore, a 

single continuous roll is able to stretch over an entire raceway45. The film attaches to a 

frame structure along the edges of the raceway so that a gas tight compartment is formed. 

Notice that the film is not taught over the top of the berms, but rather is slack. This is to 

allow for variations in hydrogen production rate which is expected with varying solar 

insolation. The number of raceways for a given reactor bed will depend on hydrogen 

production rate. Thus far it appears that the limiting factor is the width of the transparent 

roll sizes. We do not want to bond multiple sheets of LDPE as it will only provide 

additional leakage points for our product gas. 

The Pond Edging envisioned would be a mechanical fixture that would seal the transparent 

film to the pond lining as shown in Figure 3-13. 

 

44 Blake, Daniel M. “Hydrogen Reactor Development and Design for Photofermentation and Photolytic 

Processes”. NREL Project PD19. May 23-26, 2005.

45 Agriculture Films by Berry Plastics. 

http://www.covalenceplastics.com/site/content/agricultural/agricultural.asp#gh. Date Accessed 18 July 2008.
40 

Figure 3-13. Film-Frame interface point

Earthen Berms

Gas Tight Frame

Film Attachment Points

Inlet Ports: Makeup

Water & Nutrients

3.1.5.2 Paddlewheel Mixers

Two paddlewheels and a center berm are located in each raceway to provide circulation of 

the algae solution for proper mixing. The paddlewheel mixer would look similar to a water 

wheel at a mill. Our analysis for the amount of mixing necessary is based on typical mixing 

parameters for algal growth in the biodiesel industry. Using that data and resizing to our 

raceway configuration, we obtained comparable parameters which are listed in Figure 3-14. 

Paddlewheels with higher flow rates of up to 60cm/s and larger may be necessary to assist in 

the thermal control of the pond.46 The B-4 system consists of immobilized mats of algae, 

and, therefore, will not need paddlewheels for mixing. The nutrients will be provided 

through a perforated PVC tube that travels the length of the raceway to ensure adequate 

nutrient dispersal.

Figure 3-14. Paddlewheel Electrical Requirements

Biodiesel Equipment for Algal Growth

Algae Concentration 1% volume of pond

Flowrate 20 cm/s

Pond Size 1000 m2

Paddle Wheel (single) 1.23 kWh/day

Source: http://www.bioking.lv/algae_cultivated_in_raceway_ponds.html

Reactor Raceway

Algae Concentration 0.02 volume of pond

Flowrate 20 cm/s

Pond Size 4050 m2

Size Factor 4.05

Paddle Wheel (two) 2.46 kWh/day

3.1.5.3 Ports and Valves

In the Top View of Figure 3-10 one can see that the bed is made up of several raceways, 

each with an inlet and two outlet ports. At each of these ports are manual valves to isolate 

individual raceways or regulate the flow to or from each raceway. From our research it is 

possible for the daily water level drop due to evaporation to be as much as 6.2mm for a 

 

46 R. Vitale. Waterwheel Factory Inc. Telephone Interview. 08 October 2008.
41 

1,000m2 pond.47 This would translate into makeup water requirements of 28.8m3/day in our 

raceway without including the slip stream flow rates that are part of the Chemostat II and 

Chemostat reactor bed forms. In this subassembly we only consider the costs of the flanges 

and valves. All piping runs are accounted for in the subassemblies they support. 

3.1.5.4 Capital Costs

The materials selected and unit costs of the components are listed in Figure 3-15. The plant 

is anticipated to operate for 20 years. However, the transparent film is susceptible to UV 

damage which causes the film to haze and thereby become non-transparent. Thus the costs 

of that material shown below are only for the initial material. Replacements costs of the 

film are considered an operational expense. In our cost analysis (to be described) we will 

assume the polyethylene is good for 5 years. All other components of this subassembly are 

anticipated to last the entire plant life.

 

47 http://www.bioking.lv/algae_cultivated_in_raceway_ponds.html. Date Accessed 08 October 2008.
42 

Figure 3-15. Capital Costs of Reactor Bed Components 

Reactor Bed

System 

Material Chosen Unit Pricing

Transparent Film Tufflite Brand Products , available up to 56’ wide,

Polyethylene, 4-6mil

$ 0.54/m2

Source: Berry Plastics 

(http://www.covalenceplastics.com/site/content/agricultural/agricultu

ral.asp#gh) 

Pond Lining Either Butyl rubber, PVC, or LDPE (low density 

polyethylene), 30 – 40 mil

$0.47/m2

Source: 

Pond Edging Hard Plastic $ 7/m

Source: DTI Estimate

Paddlewheel Mixer 

with Motor

14 blades (6’ long x 6” wide), 3⁄4 hp motor, capable 

of 1ft/s flow

$ 5000 ea.

Source: Waterwheel Factory, Inc.

Manual Gas Valve IPEX VKD series PVC Ball Valve Part# VKDBV108, 

EPDM/Teflon Seals 2”

$ 67.23

Source: http://www.valvestore.com

Manual Water Valve IPEX VKD series PVC Ball Valve Part# VKDBV108, 

EPDM/Teflon Seals 2”

$ 43.46

Source: http://www.valvestore.com

Water Port ISO-QF flange, Schedule 80 Polyvinyl Chloride, Ref 

#PVC50-SF 

$ 8

Source: http://www.ancorp.com/line.aspx?id=819 

Gas Port ISO-QF flange, Schedule 80 Polyvinyl Chloride, Ref 

#PVC50-SF 

$ 8

Source: http://www.ancorp.com/line.aspx?id=819 

Excavation/Installation Based on 5 day excavation, 6 workers, medium 

sized equipment and no dirt removal.

$26,083/raceway

Source: Dept of Labor, Davis Bacon Wage Determinations (CA Labor 

Rates) & The Blue Book of Building & Construction

3.1.5.5 Heat Exchangers

The covered reactor area in direct sunlight could increase bed temperatures significantly. 

Heat exchangers would require large electricity and/or water requirements. In order to keep 

costs low we are assuming that the earth will provide a sufficient heat sink for this system so 

no heat exchangers would be required. 

3.2 Organism Feed Subassembly

The organism feed subassembly provides the necessary nutrients, medium and gases to 

grow and maintain a hydrogen producing algal or bacteria colony. The components of this 

subassembly are listed in Figure 3-16 and shown in Figure 3-17. The components of the 

subassembly can be further specified for a given organism and reactor bed form. A variation 

of this system will be required in the B-4 pathway, where algae growth beds and hydrogen 

production beds require different nutrients simultaneously. 
43 

Figure 3-16. Organism Feed Components

Organism Feed Subassembly 

Components

Organism Feed Subassembly 

Consumables

Nutrient/Fertilizer Tank CO2

Metering Pump Nutrients/Fertilizer

Mixing Pump Water (Initial Fill)

CO2 Cylinder

CO2 Saturation Vessel

Transfer In Pump

Piping

Figure 3-17. Organism Feed Subassembly Design

Organism Feed

Subassembly

CO2

from EtOH

plant

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

from

Recycle

Subassembly

Photo bio

Reactor Bed

Water Flow

Valve

3.2.1 Flow Valves

The flow valves are shown in this flow diagram because they are located within the area of 

the organism feed subassembly. However because they will be monitored and operated 

remotely, their costs are accounted for in the control system subassembly. For further 

description of what types of valves are selected refer to that section of the document. 
44 

3.2.2 Vessels and Tanks

There are three vessels in this subassembly. The CO2 Cylinder is a standard off the shelf 

item from one of the gas producer such as Air Products which supplies to industrial plants. 

The fertilizer is envisioned to be in aqueous form. The fertilizer will arrive at the facility 

pre-mixed for our application. The tank can be a stationary item on its own concrete pad 

that is refilled as necessary. The alternative would be a trailer provided by the fertilizer 

supplier which is dropped off when the trailer on site is dry. 

3.2.3 Pumps

This subassembly has several pumps. After further research, centrifugal pumps were 

considered the most cost effective and appropriate pumps for our application. The pumps 

are sized based on the flow rates required to operate the reactor. Thus given a fixed 

hydrogen production rate of 1,000kgH2/day, pump size will vary depending on the form of 

the reactor bed. 

3.2.4 Capital Costs

The capital costs associated with this subassembly have been identified in Figure 3-18. A 

distinction is made between the consumables needed to establish the initial colony (which 

are included in capital costs of the Organism Feed Subassembly) and consumables needed 

for subsequent operation of the production plant (which will be included as feedstock costs 

under plant operation expenses). In a Chemostat II reactor bed, we postulate that CO2 gas 

will be necessary during the initial growth stage of the algae and, thus, are included in the 

plant capital costs. The costs of CO2 required during the reduced growth stage will be 

accounted for in plant operational expenses. The fertilizer will need to be continually fed to 

the colony during hydrogen production. Transporting liquids into and out of the reactor bed 

will require piping which is considered part of this subassembly. Pump sizes are a 

representative average in Figure 3-18. Actual sizes are in the Bill of Materials in section 4. 
45 

Figure 3-18. Capital Cost of Organism Feed Components

Organism Feed

Subassembly 

Material Chosen Unit Pricing

Components

Nutrient Metering 

Pump

1468 gph $2,594

Source: Perry's Chemical Handbook, Plant Design and 

Economics/Peters and Timmerhaus

Transfer In Pump 316 SS centrifugal pump 7.5KW Motor, 

150 gpm capacity

$10,252

Source: Perry's Chemical Handbook, Plant Design and 

Economics/Peters and Timmerhaus

Mixing Pump 316 SS centrifugal pump 7.5KW Motor, 

150 gpm capacity

$10,252

Source: Perry's Chemical Handbook, Plant Design and 

Economics/Peters and Timmerhaus

CO2 Cylinder Aluminum 6061-T6 alloy, 1800 psi Service 

Pressure, 50 lb. Capacity Part #6125

$ 360

Source: http://kegman.net/carbon.htm#aluminumtank 

CO2 Saturation Vessel S.S. field erected, 14,760 gal capacity $30,000

Source: 

Nutrient Tank 10 gallon Cylindrical Process Tank, Part 

#0275-085

$ 30

Source: http://www.watertanks.com/products/0275-085.asp

Raceway Slurry 

Collection Piping

1” diameter, plastic $1.00/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Main Slurry Collection 

Piping (B1/B2)

2” diameter, plastic $2.12/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Main Slurry Collection 

Piping (B5)

3” diameter, plastic $4.31/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Main Slurry Collection 

Piping (B3)

4” diameter, plastic $6.18/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Recycle-Feed Transfer 

Piping (B1/B2)

3” diameter, plastic $4.31/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Recycle-Feed Transfer 

Piping (B5)

4” diameter, plastic $6.18/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Recycle-Feed Transfer 

Piping (B3)

5.5

” diameter, plastic

$8.51/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Consumables

Nutrient, Initial Colony 

Growth

3 lb/Acre/yr ~$ 0.49/lb

Source: USDA 

CO2 Gas 50 lb cylinder refill at distributor facility $ 35

Source: NFC Company, Chicago, IL (800) 734 4515

Water Standard Process Water, not Salt Water $0.001665/gal

Source: H2A Feedstock Costs, 2008
46 

3.3 Recycle Subassembly

The purpose of the recycle subassembly is to separate waste organisms from water so that 

the water can be reused in the reactor bed. The components of this subassembly in general 

terms are listed in Figure 3-19 and shown in Figure 3-20. 

Figure 3-19. Recycle Components

Recycle Subassembly Components

Transfer Out Pump

Algal Separation Unit

Piping

In this subassembly the organism solution is pumped to a device capable of separating the 

organism from the water. This is envisioned as a rotary drum filter. However, several other 

systems were considered. The organism-water separator will result in both organism and 

water outflows. The water outflow is recycled and combined with new water to make up for 

the amount that has evaporated, consumed in the production of hydrogen, or could not be 

separated from the organism. That water is sent back to the Organism Feed Subassembly to 

be combined with nutrients prior to being re-injected into the Photo-bio reactor bed. The 

subassembly components are all shown in Figure 3-20. 

Figure 3-20. Recycle Subassembly Design

Recycle

Subassembly

Transfer Out

Pump

Wet Spent

Makeup Algae

Water

Algal Separation

Wet Spent

Algae

Recycle

Subassembly

to

Algae Feed

Subassembly

Photo bio

Reactor Bed

to

Algae Feed

Subassembly
47 

3.3.1 Organism Separation Concepts

In order to maintain a constant volume of living organisms in the system, it will be 

necessary to remove both dead and excess living organisms stemming from the continual 

growth of the culture. It is also imperative to separate the organisms from the water to allow 

for maximum water recycling. Two different concepts for organism separation were 

examined: industrial centrifuges and rotary drum filters. In the end, rotary drum filters are 

the most economically and technically viable option, due to their lower capital cost and 

much higher rate of water return.

3.3.1.1 Centrifuges

In most of the research concerning photobiological production of hydrogen using algae or 

bacteria, a laboratory centrifuge is used to separate the organisms from the nutrient medium. 

Owing to this precedent, we initially investigated industrial sized centrifuges as a means of 

algae separation. Consultation with many centrifuge manufacturers including Alfa Laval, 

Westphalia, US Centrifuge, and Centrifuge Experts International indicated that while 

centrifuges are able to handle large flow rates, the extremely low concentration of organisms 

in the mixture would prove difficult to separate well. Capital costs for this system would 

also be fairly expensive since many machines running at flow rates far below their rated 

capacity would be necessary to effectively separate the organisms. 

After consultation, the two kinds of industrial centrifuges considered were vertically spun 

and decanter centrifuges. Conversations with representatives from the aforementioned 

companies all suggested that vertically spun centrifuges would be more adept at 

algae/bacteria separation than decanter. The decanter centrifuges generally offered higher 

flow rates, but extremely poor separation performance. Dialogue with representatives from 

Centrifuge Experts International indicated that in order to get an appreciable percentage of 

water return from a vertically spun centrifuge, the machines would have to be run at 15% of 

rated capacity, necessitating multiple machines. One potential avenue for increasing the 

flow rate while maintaining a high water recovery percentage would be through the use of 

chemical flocculants. The drawback of this method is the contamination of the water, a 

solution to which is still being developed. At nearly $400,000 each, with the necessity for a 

minimum of four machines, the capital costs for this system are fairly high, especially given 

the water recovery rate of only 90%. Westphalia centrifuges yielded similar results with 

slightly higher capital costs. Details about the various types of centrifuges can be seen in 

Figure 3-21. 

Figure 3-21. Types of Centrifuges 

Company Type of Filtration Flow Rate (GPH) Motor Required Type Cost Quantity 

needed Drawbacks

Centrifuge Experts 

International Vertically Spun 24, 000 50 HP Seital SE 10 $400, 000 2-4 Unable to run at full flow 

rate

Westphalia Vertically Spun 15, 600 75 KW HSB 400 550, 000 2-4 Unable to run at full flow 

rate

US Centrifuge Decanter 22, 500 150HP/ 25HP Model CQ 7 $400, 000 1 Not able to separate algae 

well

Alfa Laval Decanter 30, 000 250HP/ 50HP Aldec G2 $750, 000-$800, 

000 1 Not able to separate algae well

 

 

Algae Separation

Company Type of Filtration Flow Rate (GPH) Motor Required Type Cost Quantity 

needed Drawbacks

Centrifuge Experts 

International Vertically Spun 24, 000 50 HP Seital SE 10 $400, 000 2-4 Unable to run at full flow 

rate

Westphalia Vertically Spun 15, 600 75 KW HSB 400 550, 000 2-4 Unable to run at full flow 

rate

US Centrifuge Decanter 22, 500 150HP/ 25HP Model CQ 7 $400, 000 1 Not able to separate algae 

well

Alfa Laval Decanter 30, 000 250HP/ 50HP Aldec G2 $750, 000-$800, 

000 1 Not able to separate algae well

 

 

Algae Separation
48 

3.3.1.2 Rotary Drum Filters

Considering the limitations of conventional centrifuge separation, research was done into 

the aquaculture industry, which deals heavily with algae and algae separation. Discussions 

with Advanced Aquaculture Systems Inc. and information from www.phyco.org indicated 

that rotary drum filters are far superior to centrifuges for organism separation. In rotary 

drum separation the input water from the pond is pumped into the filter. The organisms

catch on the inside of a <10 micron screen in the rotating drum and the clean water passes 

through the screen as a backwash system rinses the solids of the screen into a sludge trough. 

A diagram of a drum filter can be seen in Figure 3-22. Expected water return from such a 

system is greater than 99%. The capital costs are also lower. Three Hydrotech drum filters 

will be needed to handle the flow rate of one B-1 module, but are only $87,000 per drum, 

substantially lower than the centrifuges. 

Figure 3-22. Drum Filter Diagram48

3.3.2 Pumps

This subassembly has a single pump. A centrifugal pump was considered the most cost 

effective and appropriate for our application. The pump is sized based on the flow rates 

required to operate the reactor. Thus given a fixed hydrogen production rate of 

1,000kgH2/day, pump size will be different depending on the form of the reactor bed. The 

materials selected and unit costs of the components needed for this subassembly are listed 

in Figure 3–23 

 

48 Email from Dana Kent. (Advanced Aquaculture Inc.) 19 August 2008.
49 

Figure 3-23. Capital Costs of Recycle Components

Recycle 

Subassembly 

Material Chosen Unit Pricing

Rotary Drum filter Hydrotech Drum Filter - 3750 gph $87,000 ea.

Source: Advanced Aquaculture Inc. 

Transfer Out Pump 316 SS centrifugal pump 7.5KW Motor, 

150 gpm capacity

$10,252

Source: Perry's Chemical Handbook, Plant Design and Economics/Peters 

and Timmerhaus

Main Slurry Feed Piping 

(B1/B2)

2” diameter, plastic $2.12/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

Main Slurry Feed Piping 

(B5)

3” diameter, plastic $4.31/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

Main Slurry Feed Piping 

(B3)

4” diameter, plastic $6.18/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

Raceway Slurry Feed 

Piping

1” diameter, plastic $1.00/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

3.4 Gas Capture Subassembly

The function of gas capture subassembly is to compress and separate our product hydrogen 

and deliver it to the production facility limits. The components of this subassembly in 

general terms are listed in Figure 3-24. 

Figure 3-24. Gas Capture Components

Gas Capture Subassembly Components

Compressor

Gas Separation

Piping

The outlet pressure of hydrogen at the plant gate is 300psi. This pressure was selected to 

provide a system comparable to other DOE H2A Production Plants. Those cases have been 

normalized to an outlet pressure of 300psi. The hydrogen is at atmospheric pressure when 

released by the organisms. The Hydrogen Flow Meter is located at this subassembly but it 

is accounted for in the control system subassembly. Between these two points in the 

subassembly there is minimally a compressor and a gas separation unit. The compressor 

will need to bring the gas mixture to the pressures required for separation and transport to 

the plant gate. Pressure losses are possible in the gas separation unit and the pipeline runs 

so compression may be higher than 300psi. The subassembly components are all shown in 

Figure 3-25. 
50 

Figure 3-25. Gas Capture Subassembly Design

Gas Capture

Subassembly

Hydrogen

Pipeline

Compressor

Hydrogen

Pipeline

Gas

Separation

Gas Capture Subassembly

from

Photo bio

Reactor Bed

Hydrogen

Flow Meter

3.4.1 H2-O2 Gas Mixture Safety 

Since the organisms create hydrogen via water splitting, oxygen is necessarily created as a 

byproduct of hydrogen production. For the systems that must maintain anaerobic conditions 

for hydrogen production (B3, B4, and B5), a separate reaction (i.e. respiration) occurs to 

virtually instantaneously consume the oxygen, so that O2 levels are maintained at very low 

levels. However, the B1 and B2 pathways use oxygen tolerant organisms that can carry out 

water splitting without a separate respiration reaction to absorb the O2, and thus they 

generate substantial O2, at a stoichiometric ratio with the produced hydrogen. This 2:1 

molar ratio of hydrogen to oxygen is a flammability concern and therefore precautions must 

be taken to eliminate any potential ignition source. H2 safety precautions are important to 

all the systems, however, the issue particularly affects the B-1 and B-2 pathways where a 

stoichiometric H2/O2 mixture is stored, i.e., in raceways, headspaces, and gas processing 

systems. For this analysis, the requisite safety sensors and controls are used to assure safe 

plant operation with this gas mixture. These kinds of safety measures are also routinely 

dealt with in other industrial systems using H2. 

3.4.2 Compressors

Conversations with Norwalk Compressors of Norwalk CT suggest that an oil-free, balancedopposed, 3-stage piston compressor can be used to safely compress the hydrogen or 

hydrogen/oxygen gas stream prior (or after) gas separation. Particular care must be taken 

when compressing flammable gas mixtures to avoid sparking or exceeding temperature 
51 

limits. However, compression of flammable mixtures is a routine industrial process 

conducted safely every day. 

Pursuant to the modular concept, the compression system is scaled for the gas flows of a 

1,000kgH2/day module. The compression module is sized for the average hydrogen 

production for the longest day of the year: the summer solstice (June 21st). At all other 

times, the compressor is operated at reduced capacity. Power for the compressor is based on 

mass flow rate and compression ratio49 which are both modest by industrial standards. 

Consequently, compressor power is moderate: approximately 78-123 kW depending on 

system. However, due to the low hydrogen pressure generated above the raceways (1atm), 

compressor piston diameter for the first stage will be quite large. For this reason, diaphragm 

compressors, which would normally be attractive for their longer life and non-leak 

attributes, are not viewed as viable candidates. 

Since the biohydrogen compressors are similar in type (piston compressors) and 

compression ratio (~20) to the compressors used in the DOE’s H2A forecourt analyses

50, we 

have adopted the H2A compressor cost algorithm of $4,580/ (kgH2/h). However, since this 

algorithm is based on hydrogen compression and the biohydrogen systems produce a 

mixture of gases, we convert the costing algorithm to a molar basis: $9,233/ (kgmol/h). 

3.4.3 Gas Separation Concepts

There are several commercial means available for separating hydrogen gas from a gas 

mixture. In this analysis we have focused our attention on pressure swing adsorption (PSA), 

temperature swing adsorption (TSA), membranes, electrochemical pumps and combinations 

of these methods. 

3.4.3.1 Pressure Swing Adsorption 

A pressure swing adsorption (PSA) system is preliminarily selected as the H2 separation 

system. PSA technology takes advantage of a materials affinity to preferentially adsorb a 

particular gas species at high gas pressure. PSA systems are commonly used in the 

petroleum industry to purify a variety of gases. Separation of hydrogen gas from steam 

methane reforming (SMR) product gases (H2, CO, CO2, H2O, CH4) is common. We 

propose to use PSA to separate hydrogen from a H2, O2, H2O mixture, a much less 

frequently encountered application. Dialogue with Adsorption Research, Inc. indicated that 

it was feasible to use PSA to separate such a mixture.51

PSA operate by flowing a pressurized stream of gases across a multi-component adsorbent 

bed (commonly layers of activated carbon, zeolite, silica gel, etc.) to preferentially capture 

an undesired gaseous species on the surface of the adsorbent. By careful selection of 

adsorbent materials, all undesirable species may be captured in the bed so that only high 

purity gas (often greater than 5 nines purity) exits the bed. PSA systems are inherently 

cyclic with a series of bed (typically 4, 6, 8 or 12) operating out of phase. After a bed is 

 

49 Compression ratio is 300psi/14.7psi = 20.4.

50 The standard DOE H2A forecourt analysis assumes a 4-stage, piston compressor, 300psi inlet, 6250psi 

outlet, and approximate 1500kgH2/day flow rate.

51 Knaebel, Kent. Phone conversation. 20 Oct 2008
52 

“full” (i.e. the adsorbent material no longer has open sites on which to further adsorb 

contaminant gases), the bed goes into a “vent’ cycle where pressure is decreased (typically 

to 1atm but sometimes < 1atm) to desorb the contaminants. This vent gas stream is 

typically vented to the atmosphere or fed to a burner if it contains fuel value. After venting, 

the bed is purged with pure gas stream reserve flowed i.e. fed from the exit of the bed to 

drive any residual containment gas back towards the inlet. After purging, the bed enters an 

“equalization” cycle where gas pressure is raised to full operating pressure in preparation for 

the resumption of contaminant gas flow. 

 

Hydrogen recovery is a key metric of PSA performance. Hydrogen recovery is defined as 

the fraction of inlet gas hydrogen that is captured in the pure gas stream. There are two 

main sources of hydrogen loss. The first is hydrogen that is contained in the bed at the 

beginning of the vent cycle. As the bed is depressurized, this hydrogen is expelled and lost 

to recovery. The second source of loss is during the purge cycle when pure hydrogen is 

used to actively vent the system of impurities. By summing these two losses, an accurate 

measure of hydrogen recovery can be attained. 

To determine H2 recovery for an H2/O2 mixed gas system, we have ascertained the packing 

density and adsorption performance of carbon (for O2 adsorption) and silica gel (for water 

adsorption). These parameters are listed in Figure 47. By carefully modeling the desired 

flow rates and calculating the hydrogen contained in the bed during the various cycles, the 

hydrogen recovery is calculated. 

Key parameters of the PSA system are shown in Figure 3-26. Parameters are listed for the

adsorption of water, methane, carbon dioxide, carbon monoxide, nitrogen, and oxygen. 

Although oxygen is the primary species for adsorption, the other gases are included both 

because they were used in validation of the modeling approach and may be present in low 

quantities due to the decomposition of algae. For each adsorbent there are two values of 

interest; the amount of gas that adheres to the adsorbent at high pressure, and the amount of 

gas that adheres to the absorbent at low pressure. It is the difference in adsorption levels 

between the two pressures that is of interest. Although not necessarily linear, we have 

assumed that the adsorption data varies linearly between the two data points listed.

Capital costs for the individual PSA systems are estimated based on both the performance 

model described above and on a scaling factor approach. The performance model is used to 

calculate the approximate bed size of the PSA vessels for the particular flow rates and gas 

compositions of each system. Once this PSA bed size is determined, a scaling factor 

approach is used to ratio the cost relative to a H2A model PSA system. A 0.5 scaling factor 

is assumed resulting in the following equation: 

BaseBedSize

NewBedSize NewCost = BaseCost *

The base PSA system cost assumed is based on an H2A 1500kgH2/day SMR PSA having 

6,065L of total absorbent and a $100,000 total price.
53 

Figure 3-26. Possible PSA Parameters

Silica Gel (for Water Absorption)

Bed Void Fraction % 36%

Bed Usage Fraction (1/LUB) % 50%

Adsorp. Fraction at Specified Pressure g H20/ g adsorbent 30%

Specified Pressure for Adsorbent Data psi 250

Adsorp. Fraction at Specified Purge Pressure g H20/ g adsorbent 0.2

Specified Purge Pressure psi 14.7

Estimated adsorbent O2 Update Fraction g H2O/ g adsorbent 0.12039949

Bulk Density g/L 500

Carbon (for Methane)

Bed Void Fraction % 0.36

Bed Usage Fraction (1/LUB) % 0.8

Update Fraction at Specified Pressure g CH4/ g adsorbent 0.0448

Specified Pressure for Adsorbent Data psi 250

Adsorp. Fraction at Specified Purge Pressure g CH4/ g adsorbent 0.016

Specified Purge Pressure psi 14.7

Estimated adsorbent O2 Update Fraction g CH4/ g adsorbent 0.034675053

Bulk Density g/L 515

Carbon (for CO2)

Bed Void Fraction % 0.36

Bed Usage Fraction (1/LUB) % 0.714285714

Update Fraction at Specified Pressure g CO2/ g adsorbent 0.1694

Specified Pressure for Adsorbent Data psi 250

Adsorp. Fraction at Specified Purge Pressure g CO2/ g adsorbent 0.044

Specified Purge Pressure psi 14.7

Estimated adsorbent O2 Update Fraction g CO2/ g adsorbent 0.15098096

Bulk Density g/L 515

Zeolite (for CO)

Bed Void Fraction % 0.36

Bed Usage Fraction (1/LUB) % 0.714285714

Update Fraction at Specified Pressure g CO/ g adsorbent 0.042

Specified Pressure for Adsorbent Data psi 250

Adsorp. Fraction at Specified Purge Pressure g CO/ g adsorbent 0.0364

Specified Purge Pressure psi 14.7

Estimated adsorbent O2 Update Fraction g CO/ g adsorbent 0.006742371

Bulk Density g/L 515

Zeolite (for N2)

Bed Void Fraction % 0.36

Bed Usage Fraction (1/LUB) % 0.3

Update Fraction at Specified Pressure g N2/ g adsorbent 0.04725

Specified Pressure for Adsorbent Data psi 250

Adsorp. Fraction at Specified Purge Pressure g N2/ g adsorbent 0.0406

Specified Purge Pressure psi 14.7

Estimated adsorbent O2 Update Fraction g N2/ g adsorbent 0.008006566

Bulk Density g/L 515

Zeolite (for O2)

Bed Void Fraction % 0.36

Bed Usage Fraction (1/LUB) % 0.769230769

Update Fraction at Specified Pressure g O2/ g adsorbent 0.032

Specified Pressure for Adsorbent Data psi 250

Adsorp. Fraction at Specified Purge Pressure g O2/ g adsorbent 0

Specified Purge Pressure psi 0.001

Estimated adsorbent O2 Update Fraction g O2/ g adsorbent 0.036262545

Bulk Density g/L 680

 Where:

 Bed Void Fraction = interstitial void fraction between adsorbent particles

 Bed Usage Fraction = 1/(Length of Unused Bed) = Fraction of bed used for adsorbent calcs 

where the remained is buffer against contaminant breakthrough.

3.4.3.2 Temperature Swing Adsorption

Temperature swing adsorption (TSA) systems are analogous to PSA systems except they 

use differences in temperature rather than pressure to preferentially adsorb and desorb 

contaminant species. While used industrially, TSA systems are not as prevalent in process 

plants as PSA systems. The main advantage of a TSA system is that it would not require 

compression of the H2-rich gas thereby avoiding the safety concerns of compressing H2/O2
54 

gas mixtures and potentially lowering electrical requirements. However, separation of H2

and O2 via TSA requires a significant temperature swing with a refrigeration system 

required to achieve the lower temperature bound. Was “waste” cooling or heating available 

from an adjacent process, TSA would be an attractive option. However, as currently 

configured such “waste” thermal energy is not readily available and thus the refrigeration 

system would be a substantial energy and cost element. Additionally, conversations with 

Kent Knaebel from Adsorption Research Incorporated, a gas separation consulting 

company, indicated that PSA was a far superior option, given its vastly reduced cycle time. 

The cycle for a TSA system would be measured in hours, as opposed to minutes for PSA, 

necessitating a much larger bed size in order to separate the same amount of gas. This 

increased capital cost of the beds and cooling system makes TSA less economically 

practical than PSA. Consequently, we have not pursued further TSA system analysis for 

any of the bio-hydrogen systems.

3.4.3.3 Membrane Separation

Membrane separation units are of two main types: metallic membranes and nano-porous 

membranes. Metallic membranes typically use palladium (Pd), Pd-alloys, or layered 

Pd/alloy to allow the diffusion of H+ ions through the membrane. Pd membranes are 

characterized by very high hydrogen selectivity (typically <10,000), high cost (due to the 

inherent Pd material cost), and moderate hydrogen permeability that is primarily a function 

of temperature and membrane thickness. Hydrogen flux through the membrane follows 

Boyds Law and is proportional to the difference of the square root of the reformate 

(upstream) hydrogen partial pressure and the square root of the permeate (downstream) 

hydrogen partial pressure. Consequently, Pd-membrane systems work best with highly 

pressurized inlet streams and low pressure H2 product streams. While a few small scale Pdmembrane commercial products are on the market (e.g. IdaTech), large scale Pd-membrane 

hydrogen purification systems are not employed industrially due to performance and cost 

concerns. Based on membrane system modeling for a bio-hydrogen system with 60% H2 at 

20atm (300psi) and a 1atm permeate (pure H2) outlet pressure, hydrogen recoveries in 

excess of 90% are theoretically possible. 

Drawbacks of Pd-membrane separation systems include: 

• the necessity to heat the membrane (and hydrogen) to 250-350o

C to ensure adequate 

hydrogen permeability 

• the requirement to compress the H2-rich feedstream to high pressure thereby raising 

safety concerns related to H2/O2 compression

• the low pressure of the pure H2 product stream thereby requiring additional H2

compression to achieve a pipeline transport pressure 

• uncertainty of using Pd membrane with H2/O2 gas mixtures (the Pd may oxidize or 

be an H2/O2 combustion catalyst) 

• the general immaturity of the technology. 

For these reasons, Pd-membranes are not selected for further analysis.

The other broad class of membrane separation systems is based on nano-porous materials. 

Nano-porous materials function as molecular sieves and use pore size to preferentially pass 
55 

molecular hydrogen. As reproduced in Figure 3-27, Phair and Badwal52 have surveyed the 

range of nano-porous membrane options. Of these, we judge polymeric membrane to be 

most applicable to bio-hydrogen gas separation due to their low temperature of operation. 

However, polymer membranes are not highly selective and thus would require additional 

downstream purification. Additionally, hydrogen flow is driven by differences in hydrogen 

partial pressure across the membrane. Consequently, compression of the unseparated gas is 

required once again raising safety concern for H2/O2 gas mixtures. A detailed analysis 

beyond the scope of this project is required to optimize pressure level, permeate purity, and 

recover H2. For these reasons, polymer membranes don’t appear to offer substantially 

superior benefits over PSA systems. Consequently, polymer membranes are not selected for 

further analysis. 

Figure 3-27. Nano-porous Membranes

3.4.3.4 Electrochemical Pumps 

Electrochemical purification of hydrogen is possible using an applied voltage to drive 

hydrogen across a separation membrane. Such systems have been demonstrated at small 

scale but are not currently in industrial use. In addition to the gas separation function, 

electrochemical pumps can be used to pressure the hydrogen stream thereby eliminating or 

reducing the need for mechanical H2 gas compression. In theory, a mixed gas stream could 

enter the electrochemical pump at 1atm and a high H2 purity gas stream could exit at 

pressure (10-100atm). 

Conversations with Glen Eiseman of Hydrogen Pump LLC preliminarily explored the use of 

electrochemical pumps for biohydrogen purification. Current products from Hydrogen 

Pumps use a PBI membrane, operate at ~160o

C, employ a Pt catalyst, compress the H2

stream to several atmospheres, and contain a supplementary gas cleanup system since only 

98% pure H2 is passed across the membrane. Since exposure of a H2/O2 gas mixture to Pt 

would result in catalyst combustion of the gases, such a combination is unacceptable. 

Development of alternate non-Pt catalyst is theoretically possible, but the authors know of 

no such development efforts. 

 

52“Materials for separation membranes in hydrogen and oxygen production and future power generation”, J.W. 

Phair, S.P.S. Badwal, (CSIRO Manufacturing and Infrastructure Technology, Victoria, Australia) Science and 

Technology of Advanced Materials 7 (2006) 792–805. 
56 

Electrochemical pumps could be used for separation of non-O2 containing gas mixtures. 

However, the systems would incur the following disadvantages: 

• the necessity to heat the membrane (and hydrogen) to 160oC 

• the need for a secondary gas cleanup system (such as PSA) to further purify the H2

product gas to five 9’s purity 

• the electrical consumption to power the device 

For these reasons, electrochemical pumps are not selected for further analysis.

3.4.4 Capital Costs

The components of the subassembly can be further specified for a given organism and 

reactor bed form. However we have included in Figure 3-28 materials selected and unit 

costs of components. 

Figure 3-28. Capital Cost of Gas Capture Components

Gas Capture 

Subassembly 

Material Chosen Unit Pricing

Compressor 142kgH2/hr size for 1000kgH2/day plant $9233/kgmol/hr

Source: Using H2A Cost guidelines

Gas Separation Pressure Swing Adsorption $30,542 - $119,408

Source: Using H2A Cost guidelines and scaling factors

Gas Capture Collection 

Piping

5.5” diameter, plastic, 100ft/s velocity, $8.51/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Gas Capture Collection 

Piping

4.5” diameter, plastic, 100ft/s velocity, $6.18/ft

Source: PVC Plastic Corp (http://www.usplastic.com/) 

Main Collection Piping 4” diameter, plastic, 100ft/s velocity, $6.18/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

Main Collection Piping 3” diameter, plastic, 100ft/s velocity, $4.31/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

Raceway Collection 

Piping

1.5” diameter, plastic, 100ft/s velocity, $1.57/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

Raceway Collection 

Piping

1” diameter, plastic, 100ft/s velocity, $1.00/ft

Source: PVC Plastic Corp (http://www.usplastic.com/)

3.4.5 Piping

Transporting gases out of the reactor bed and the product hydrogen to the facility limits will 

require several hundred feet of piping which are included as part of the subassemblies cost. 

For the B-3 and B-4 systems, the piping lengths needed must be doubled to account for the 

piping needed to bring CO2 to the CO2 saturation vessel. There is no drum filter for these 

systems. 
57 

3.4.6 Additional Gas Capture Subassembly Components 

It is possible for the product gas to have water saturation. In this case, a dryer may be 

needed to remove water from the gas mixture. A slight vacuum might be required to draw 

the gas mixture out of the pond network and into the gas capture subassembly. These 

components are not part of the system yet but as we finalize our analysis we’ll be able to 

determine if these components are necessary. 

3.5 Control System Subassembly

Plant control systems serve many functions including local and remote monitoring, alarming

and controlling of plant equipment and functions. The more functions the system performs, 

the more costly and complex the system becomes. A very simple system may provide only 

local indication or monitoring of equipment operation. A complex system would include all 

three functions for each piece of equipment from a remote facility and some logic for how to 

operate each piece under a given set of conditions. Automating plant operations can 

increase the cost of the control system. These costs need to be evaluated against the 

operational labor savings.

3.5.1 Components

For this subassembly the design approach was to create a very basic system which is 

practical enough to provide some savings in operational labor. Because of the reactor bed’s 

large area requirements remote capabilities are essential. This will require electrical conduit 

and wiring to be laid out between the plant components and the primary control area. 

Monitoring of only the most primary indicators will be included. Alarming capabilities of 

only the most hazardous conditions will be required. Control will only be available for flow 

valves. No automation is assumed thus a PLC will not be included. For the basic control 

system described the components and instrumentation for each function are listed in Figure 

3-29. 
58 

Figure 3-29. Control System Components 

Control System Subassembly 

Components

Control Room

Control Room Wiring Panel

Raceway Wiring Panel

Computer, Monitor, Labview software

Electrical Conduit (power and instrumentation)

Instrumentation Wiring

Instrumentation 

Monitor Alarm Control

Level Indicators Pressure Sensors Nutrient Flow Valve

Hydrogen Area Sensors Water Flow Valve

Air Temperature Indicator Oxygen Area Sensors

Water Temperature Indicator

Hydrogen Flow Meter

pH level indicator

Consumables

Electricity

While the control system is relatively simple, the subassembly is complicated by the sheer 

size of the reactor beds and the modular nature of the raceways. Each raceway must be 

individually monitored and controlled, leading to a large number of replicated sensors. 

Additionally, the instruments are dispersed over this area, ~20-66 acres in size, leading to 

very long runs of wiring and conduit. Signal boosting may be required. To simplify the 

system all instrumentation will operate on electricity rather than instrument air, so that an air 

compressor is not required. There is no power transformer considered here, rather 

electricity will be treated as a plant operating expense and bought off the grid at the typical 

$/kWh for the region. 

3.5.2 Wiring and Conduit

Because of the large size of our plant there is significant amount of wiring involved with 

conveying power to each raceway for the paddlewheels and instrumentation as well as 

bringing back signal wiring for instrumentation. Figure 3-30 shows a possible layout of the 

production facility where growth and production are carried out in a single bed. Some key

dimensions are identified so that approximate lengths of wiring and conduit can be 

computed. The number of wires required to each raceway panel is shown in the Raceway 

Wiring Panel. All panels are assumed to be at the near end of each raceway and 

subassembly center for conceptual design calculations. The hydrogen flow meter has a 

single signal and power wire. The flow valves have 1 power wire but need 2 signal wires; 1 

for actuating valve and 1 for status of valve. The remaining equipment at each subassembly 

has power wires but no signal wires. Since this is a simple control system other equipment 

(pumps, compressor, and algal separation unit) will only have local control and monitoring. 

All monitoring and sensor instrumentation has a single power and a single signal wire with 

the exception of the hydrogen area sensor. It requires two signal wires for monitoring 
59 

hydrogen and alarming at certain conditions. Water flow control into and out of each 

individual raceway will be done with manual valves so they do not factor into the control 

system. The additional power wires are for the paddlewheel motors. There is no signal 

indicating the operation of these motors to the control room. Based on the other signals the 

operator should be able to infer the operating status of the paddlewheels. 

Figure 3-30. Control System Subassembly Design

R

A

C

E

W

A

Y

9

Control Room

R

A

C

E

W

A

Y

2

R

A

C

E

W

A

Y

3

R

A

C

E

W

A

Y

4

R

A

C

E

W

A

Y

5

R

A

C

E

W

A

Y

1

R

A

C

E

W

A

Y

6

R

A

C

E

W

A

Y

7

R

A

C

E

W

A

Y

8

R

A

C

E

W

A

Y

18

R

A

C

E

W

A

Y

11

R

A

C

E

W

A

Y

12

R

A

C

E

W

A

Y

13

R

A

C

E

W

A

Y

14

R

A

C

E

W

A

Y

10

R

A

C

E

W

A

Y

15

R

A

C

E

W

A

Y

16

R

A

C

E

W

A

Y

17

X2 =100' Y2 = 1000'

Recycle Subassembly

Water Flow Valve

Algae Feed Subassembly

Nutrient Flow Valve

Gas Capture Subassembly

Hydrogen Flow Meter

Y1 = 100'

Wiring Panel

X1 = 60'

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Instrumentation

(x 18 Raceways)

Raceway Wiring Panel

Paddlewheel Motors

(x 18 Raceways)

R

A

C

E

W

A

Y

10

R

A

C

E

W

A

Y

20

For the B-1 and B-2 systems, a plant layout like this yields the lengths of conduit and wiring 

shown in Figure 3-31. Signal and power wiring will be run in separate conduits. This is 

done to ensure that there is no interference in the signal data from voltage associated with 

the power wiring. In the diagram above one can see there are two main horizontal conduits 

runs (above and below the subassemblies). That route will be used for both power and 

signal wiring. Each individual raceway will have its own conduit run along the length of the 

raceway for the raceway wiring panel located at the near end of the raceway. Each 

subassembly will have its own conduit run off the main horizontal conduit lines for its 

wiring.
60 

Figure 3-31. Control System Wiring and Conduit Quantities for B-1, B-2 

Components Signal Wiring Qty 

(ft)

Power Wiring Qty 

(ft)

Conduit Qty (ft)

Raceways 67,200 84,000 4,560

Organism Feed 

Subassembly

900 450 100

Recycle 

Subassembly

540 270 100

Gas Capture 

Subassembly

630 630 100

Totals 69,270 85,350 4,860

The raceways category in Figure 3-31 above is a summation of all the wiring required for 

each of the 20 raceway wiring panels needed in the B-1 and B-2 systems. The flow valve 

and flow meter wiring requirements are included in the appropriate subassembly line. The 

additional power wiring in those subassembly lines are for the other equipment (pumps, 

compressors, etc) that require power but are not monitored and operated remotely. The 

control system wiring could have been done overhead with cable trays, however, since some 

level of excavation will already take place to create raceways, this analysis assumes 

underground lines. The installation costs of buried conduit in this case will be lower because 

it is included in the excavation of the raceways. 

3.5.3 Capital Costs

The analysis above provides a fairly extensive bill of materials for a typical control system. 

All of the instrumentation, wiring and conduit is assumed to be commercially on the shelf 

and readily available. The costs of those components are summarized in Figure 3-32. 

Figure 3-32. Capital Costs of Control System Components 

Control System 

Subassembly

Material Chosen Unit Pricing

Control Room 8’ x 20’ trailer $50/ft2

Source: 

http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html 

Control Room 

Wiring Panel

Customized to wire count $3000

Source: Consultation with Innomation Systems Inc.

Raceway 

Wiring Panel

Termination panel. Features transmit and, 

adjustable input/output levels. Available as PC 

board only or in a case. +6 to 15 mA Tx.

$146

Source: Information from Tessco Technologies Inc.

Computer & 

Monitor

Standard Desktop, Vostro 420 or similar $1500

Source: Dell 

Labview 

Software

Labview 8.6 Professional Edition $4299

Source: National Instruments
61 

Control System 

Subassembly

Material Chosen Unit Pricing

Level Indicators Flanged, Float level, 316 SST components, 

displays, 4 to 20 ma output, standard process 

connection

$714

Source: Omega Engineering Inc.

Pressure 

Sensors 

Solid State Pressure Transducer and meter, 316 

SST components, displays, 4 to 20 ma output, 

standard process connection

$345

Source: Omega Engineering Inc.

Hydrogen Area 

Sensors 

Thermal Conductivity Gas Analyzer, 316 SST 

components, displays, 4 to 20 ma output, 

standard process connection

$7,600

Source: Lesman Instrument Company

Air 

Temperature 

Indicator

Thermistor, 316 SST components, 4 to 20 ma 

output, standard process connection

$38

Source: Omega Engineering Inc.

Air 

Temperature 

Meter 

Panel Meter, single width, up to 24 inputs $599

Source: Omega Engineering Inc.

Water 

Temperature 

Indicator

Comes with pH sensor $0

Source: Information from Emerson Process Management

pH level 

Indicator

CPVC components, displays, 4 to 20 ma output, 

standard process connection

$435

Source: Omega Engineering Inc.

Oxygen Area 

Sensors 

Comes with Hydrogen Sensor $0

Source: Omega Engineering Inc.

Nutrient Flow 

Valve

6” vortex type, 316 SST components, displays, 4 

to 20 ma output, standard process connection

$5,500

Source: Information from Emerson Process Management

Water Flow 

Valve

6” vortex type, 316 SST components, displays, 4 

to 20 ma output, standard process connection

$5,500

Source: Information from Emerson Process Management

Hydrogen Flow 

Meter 

6” vortex type, 316 SST components, displays, 4 

to 20 ma output, standard process connection

$5,500

Source: Information from Emerson Process Management

Instrument 

Wiring 

22 GA Copper wire UL1007/UL1569 $0.02/ft

Source: Waytek Inc. at www.waytekwire.com @ 500’ qty

Power Wiring 22 GA Copper wire UL1007/UL1569 $0.02/ft

Source: Waytek Inc. at www.waytekwire.com @ 500’ qty

Conduit 1⁄2”, 100’ pack, Highly flexible black PVC tubing, 

resistant to oil, water, and corrosion

$0.58/ft

Source: Waytek Inc. at www.waytekwire.com @ 550’ qty
62 

4.

While most of the photobiological work has been done in a laboratory scaled environment 

under conditions uneconomical for large-scale production, the food industry has some 

experience at growing these organisms at larger scales. Looking at both sources, we have 

determined that the most practical means of hydrogen production with a photobiological 

system for each of the 5 photobiological pathways. 

Bill of Materials for Pathways

4.1 B-1 Pathway Production Plant

The B-1 Pathway is a Chemostat II photobiological reactor system using a mutant-antenna 

Chlamydomonas reinhardtii oxygen-tolerant variant as the H2 production microorganism. 

As a Chemostat II, the initial algal colony is grown at the normal cell growth rate until final 

colony concentration is reached, and then the process is switched over to steady-state mode 

where simultaneous growth and hydrogen production occur in the reactor beds. This is 

achieved by manipulating the nutrients and carbon dioxide provided. H2 production is 

projected at 9.2% solar-to-hydrogen efficiency and corresponds to a future optimized 

organism. Sufficient cellular activity to keep the micro-organism during H2 production 

healthy is assumed to occur with new cell growth taking place at a reduced rate compared to 

wild-type organisms (20% per day vs. ~80% per day). 

Twenty racetrack raceways of approximately 40ft width by 1090ft length are used to 

produce an average of 1,000kgH2/day throughout the year. The raceways have a liquid 

depth of 10cm to correspond to full photon capture at a cell concentration of 0.2g/L. 

Paddlewheels circulate the water/Chlamydomonas reinhardtii slurry. A slip stream of 

water/micro-organisms slurry is drawn off continuously to maintain the 0.2g/L microorganism concentration. Micro-organisms are removed from the recycle stream by a rotary 

drum filter. The slip stream volume removed is carefully matched to the growth rate to 

maintain a constant organism mass within the system (assuming a constant concentration of 

algae within the bed). The removed micro-organisms are taken to a land-fill, used in a 

subsequent fermentor, or otherwise disposed. Nutrients in the form of commercial grade 

fertilizer are mixed with the returning water to provide required nourishment to the microorganisms. CO2 is bubbled into a reservoir tank of the returning water to dissolve CO2 into 

the water for use by the micro-organisms. Only a very low amount of CO2 is needed to 

maintain cell health; thus CO2 does not appreciably accumulate in the headspace of the 

reactor beds. 

The Chlamydomonas reinhardtii micro-organism produces a net product gas of 33.33% O2

and 66.66% H2 (plus water vapor). A piston compressor is used to compress the gas 

mixture to 300psi prior to separation. A pressure swing adsorption (PSA) system is used to 

purify the hydrogen gas stream. 

Further details of the B-1 pathway are specified in Figure 2-1. The subassemblies of this 

pathway are shown in Figure 4-1. The complete bill of materials and capital costs of this 

production plant are shown in Figure 4-2. The total system cost is $2,164,488. 
63 

Figure 4-1. Production Plant Design for Oxygen-Tolerant Hydrogenase (Chlamy)

Photo Bio Reactor Bed

Mixers

Power

Source

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Control System

Subassembly

Control Room PC

Instrumentation

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Raceway Wiring Panel

Instrumentation

(x 20 Raceways)

Paddlewheel Motors

(x 20 Raceways)

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Transfer Out

Pump

Wet Spent

Organisms

Makeup

Water

Algal Separation

Recycle

Subassembly

CO2

Cylinder

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve
64 

Figure 4-2. Bill of Materials for B-1 Pathway 

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 87,080 m2 1 m2 $ 0.54 87,080 $ 46,871 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 93,998 m2 1 m2 $ 0.47 93,998 $ 44,412 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 4,000ft2.

Pond Edging 13,982 m 1 m $ 7.00 13,982 $ 97,876 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 20 $ 521,660 Based on California labor rates.

Paddlewheel Mixers 1 each $ 5,000.00 40 $ 200,000 Est. paddlewheel cost. 4 per raceway. 18 raceways.

Inlet Water Valve 0.5 in 1 each $ 43.46 20 $ 869 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 20 $ 869 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 20 $ 1,345 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 60 $ 480 Price from http://www.ancorp.com/line.aspx?id=819

Gas Capture Subassembly

Compressor 66 kgmol/hr 66 kgmol/hr $ 9,233.00 1 606,207 $ Using H2A Unit cost

PSA $ 119,407.39 1 119,407 $ Using H2A scaling

Raceway Collection Pipe 40 ft 1 ft $ 1.57 40 $ 63 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Main Collection Pipe 1080 ft 1 ft $ 6.18 1080 6,674 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Gas Capture Pipe 50 ft 1 ft $ 6.18 50 $ 309 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) 

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 2 20,503 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Nutrient Metering Pumps 1468 gph 1 each $ 2,594 1 2,594 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics. Pumps nutrients into return flow

Mixing Pump 150 gpm 150 gpm $ 10,252 2 20,503 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

CO2 Cylinder 50 lb 50 lb $ 360 1 $ 360 Aluminum 6061-T6 alloy, 1800 psi Service Pressure, 50 lb. Capacity Part #6125 (http://kegman.net/carbon.htm)

CO2 Saturation Vessel 15000 gal 15000 gal $ 30,000 1 30,000 $ Field erected Stainless Steel tank from Perry's chemical Handbook, Plant Design and Economics

Nutrient Tank 8 gal 10 gal $ 30.00 1 $ 30 10 gallon Cylindrical Process Tank, Part #0275-085 (http://www.watertanks.com/products/0275-085.asp)

Raceway Slurry Collection Pipe 40 ft 1 ft $ 1.00 40 $ 40 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 1080 ft 1 ft $ 2.12 1080 2,290 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle-Feed Transfer Pipe 180 ft 1 ft $ 4.31 180 $ 776 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle Subassembly

Rotary Drum filter 9639 gph 3750 gph $ 87,000.00 3 261,000 $ Information from Dana Kent at Advanced Aquaculture Inc. on Hydrotech Drum Filter

Transfer Out Pump 150 gpm 150 gpm $ 10,252 2 20,503 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Main Slurry Feed Pipe 1080 ft 1 ft $ 2.12 1080 2,290 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Raceway Slurry Feed Pipe 40 ft 1 ft $ 1.00 40 $ 40 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Consumables

Initial CO2 5241 lb 50 lb $ 35.00 105 3,675 $ Cost based on average fertilizer use for Aquaculture and average fertilizer costs from the USDA

Initial Nutrients 0.42 lb 1 lb $ 0.20 0 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 2,142,278 gal 1 gal $ 2,142,278 0.0017 $ 3,567 H2A Feedstock Costs

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 8,000 $ comes from price quote of $50/ft2 from http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 3,000.00 $ Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 20 $ 2,920 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 1,500.00 $ Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 4,299.00 $ Information from Emerson Process Management

Level Indicators 1 $ 714.00 20 14,280.00 $ Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 20 6,900.00 $ Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 10 76,000.00 $ Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/analyt_hwhydrogengas.htm)

Air Temperature Meter 1 $ 599.00 1 599.00 $ Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 20 760.00 $ Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 20 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 20 8,700.00 $ Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 20 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Instrument Wiring 69270 ft 1 ft $ 0.02 69270 1,343.84 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 85350 ft 1 ft $ 0.02 85350 1,655.79 $ Information from Waytek Inc. at waytekwire.com

Conduit 4860 ft 1 ft $ 0.58 4860 2,817.83 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 2,164,488

4.2 B-2 Pathway Production Plant

The B-2 pathway is a Chemostat II photobiological reactor system using a mutant-antenna 

cyanobacteria (often called blue-green algae) oxygen-tolerant variant as the H2 production 

microorganism. As a Chemostat II, the initial bacteria colony is grown at the normal cell 

growth rate until final colony concentration is reached, and then the process is switched over 

to steady-state mode where simultaneous growth and hydrogen production occur in the 

reactor beds. This is achieved by manipulating the nutrients and carbon dioxide provided. 

H2 production is projected at 9.2% solar-to-hydrogen efficiency and corresponds to a future 

optimized organism. Sufficient cellular activity to keep the micro-organism healthy during 

H2 production is assumed to occur with new cell growth taking place at a reduced rate 

compared to wild-type organisms (20% per day vs. ~80% per day). The B-2 pathway is very 

similar to the B-1 pathway with the primary difference being type of micro-organism. 

Twenty racetrack raceways of approximately 40ft width by 1090ft length are used to 

produce an average of 1,000kgH2/day over the year. The raceways have a liquid depth of 

10cm to correspond to full photon capture at a cell concentration of 0.2g/L. Paddlewheels 

circulate the water/bacteria slurry. A slip stream of water/micro-organisms slurry is drawn 

off continuously to maintain the 0.2g/L micro-organism concentration. Micro-organisms 
65 

are removed from the recycle stream by a rotary drum filter. The slip stream volume 

removed is carefully matched to the growth rate to maintain a constant organism mass 

within the system (assuming a constant concentration of algae within the bed). The 

removed micro-organisms are taken to a land-fill, used in a subsequent fermentor, or 

otherwise disposed. Nutrients in the form of commercial grade fertilizer are mixed with the 

returning water to provide required nourishment to the micro-organisms. CO2 is bubbled 

into a reservoir tank of the returning water to dissolve CO2 into the water for use by the 

micro-organisms. Only a very low amount of CO2 is needed to maintain cell health; thus 

CO2 does not appreciably accumulate in the headspace of the reactor beds.

 

The cyanobacteria micro-organisms produce a net product gas of 33.33% O2 and 66.66% H2

(plus water vapor). A piston compressor is used to compress the gas mixture to 300psi prior 

to separation. A pressure swing adsorption (PSA) system is used to purify the hydrogen gas 

stream.

Further details of the B-2 pathway are specified in Figure 2-1. The subassemblies of this 

pathway are shown in Figure 4-3. The complete bill of materials and capital costs of this 

production plant are shown in Figure 4-4. The total system cost is $2,164,488. 

Figure 4-3. Production Plant Design for Oxygen-Tolerant Hydrogenase 

(Cyanobacteria) 

Photo Bio Reactor Bed

Mixers

Power

Source

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Control System

Subassembly

Control Room PC

Instrumentation

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Raceway Wiring Panel

Instrumentation

(x 20 Raceways)

Paddlewheel Motors

(x 20 Raceways)

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Transfer Out

Pump

Wet Spent

Organisms

Makeup

Water

Algal Separation

Recycle

Subassembly

CO2

Cylinder

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve
66 

Figure 4-4. Bill of Materials for B-2 Pathway 

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 87,080 m2 1 m2 $ 0.54 87,080 $ 46,871 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 93,998 m2 1 m2 $ 0.47 93,998 $ 44,412 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 4,000ft2.

Pond Edging 13,982 m 1 m $ 7.00 13,982 $ 97,876 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 20 $ 521,660 Based on California labor rates.

Paddlewheel Mixers 1 each $ 5,000.00 40 $ 200,000 Est. paddlewheel cost. 4 per raceway. 18 raceways.

Inlet Water Valve 0.5 in 1 each $ 43.46 20 $ 869 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 20 $ 869 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 20 $ 1,345 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 60 $ 480 Price from http://www.ancorp.com/line.aspx?id=819

Gas Capture Subassembly

Compressor 66 kgmol/hr 66 kgmol/hr $ 9,233.00 1 606,207 $ Using H2A Unit cost

PSA $ 119,407.39 1 119,407 $ Using H2A scaling

Raceway Collection Pipe 40 ft 1 ft $ 1.57 40 $ 63 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Main Collection Pipe 1080 ft 1 ft $ 6.18 1080 6,674 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Gas Capture Pipe 50 ft 1 ft $ 6.18 50 $ 309 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) 

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 2 20,503 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Nutrient Metering Pumps 1468 gph 1 each $ 2,594 1 2,594 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics. Pumps nutrients into return flow

Mixing Pump 150 gpm 150 gpm $ 10,252 2 20,503 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

CO2 Cylinder 50 lb 50 lb $ 360 1 $ 360 Aluminum 6061-T6 alloy, 1800 psi Service Pressure, 50 lb. Capacity Part #6125 (http://kegman.net/carbon.htm)

CO2 Saturation Vessel 15000 gal 15000 gal $ 30,000 1 30,000 $ Field erected Stainless Steel tank from Perry's chemical Handbook, Plant Design and Economics

Nutrient Tank 8 gal 10 gal $ 30.00 1 $ 30 10 gallon Cylindrical Process Tank, Part #0275-085 (http://www.watertanks.com/products/0275-085.asp)

Raceway Slurry Collection Pipe 40 ft 1 ft $ 1.00 40 $ 40 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 1080 ft 1 ft $ 2.12 1080 2,290 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle-Feed Transfer Pipe 180 ft 1 ft $ 4.31 180 $ 776 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle Subassembly

Rotary Drum filter 9639 gph 3750 gph $ 87,000.00 3 261,000 $ Information from Dana Kent at Advanced Aquaculture Inc. on Hydrotech Drum Filter

Transfer Out Pump 150 gpm 150 gpm $ 10,252 2 20,503 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Main Slurry Feed Pipe 1080 ft 1 ft $ 2.12 1080 2,290 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Raceway Slurry Feed Pipe 40 ft 1 ft $ 1.00 40 $ 40 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Consumables

Initial CO2 5241 lb 50 lb $ 35.00 105 3,675 $ Cost based on average fertilizer use for Aquaculture and average fertilizer costs from the USDA

Initial Nutrients 0.42 lb 1 lb $ 0.20 0 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 2,142,278 gal 1 gal $ 2,142,278 0.0017 $ 3,567 H2A Feedstock Costs

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 8,000 $ comes from price quote of $50/ft2 from http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 3,000.00 $ Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 20 $ 2,920 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 1,500.00 $ Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 4,299.00 $ Information from Emerson Process Management

Level Indicators 1 $ 714.00 20 14,280.00 $ Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 20 6,900.00 $ Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 10 76,000.00 $ Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/analyt_hwhydrogengas.htm)

Air Temperature Meter 1 $ 599.00 1 599.00 $ Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 20 760.00 $ Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 20 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 20 8,700.00 $ Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 20 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Instrument Wiring 69270 ft 1 ft $ 0.02 69270 1,343.84 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 85350 ft 1 ft $ 0.02 85350 1,655.79 $ Information from Waytek Inc. at waytekwire.com

Conduit 4860 ft 1 ft $ 0.58 4860 2,817.83 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 2,164,488

4.3 B-3 Pathway Production Plant

The B-3 pathway is a single-bed photobiological reactor system using a mutant-antenna 

Chlamydomonas reinhardtii sulfate permease variant as the H2 production micro-organism. 

As a single bed reactor, the initial algal colony is grown at the normal cell growth rate until 

final colony concentration is reached. After sealing the system, the oxygen evolution rate, 

reduced due to the mutation in the sulfate permease proteins, falls below the rate of 

oxidative respiration, leading to an anaerobic system that produces hydrogen for a period of 

roughly 3 days. When the organism’s capability to produce hydrogen under these 

conditions has diminished, the system undergoes a 4-day regeneration/growth phase 

wherein O2 is given to the micro-organisms, halting H2 production through hydrogenase 

inhibition and encouraging cellular growth. This is followed by another 3-day H2

production phase under anaerobic conditions, when the outside O2 addition is stopped. H2

production is projected at 5.2% average solar-to-hydrogen efficiency and corresponds to a 

future optimized organism. 

Thirty-eight raceways of approximately 40ft width by 1090ft length are used to produce an 

average of 1,000kgH2/day throughout the year. The raceways have a liquid depth of 10cm 

to correspond to full photon capture at a cell concentration of 0.2g/L. In order to allow 
67 

enough growth during the regeneration phase, however, the concentration will be higher 

during reactor operation. Paddlewheels circulate the water/Chlamydomonas reinhardtii

slurry. Nutrients in the form of commercial grade fertilizer are mixed with make-up water 

stream to provide required nourishment to the microorganisms. 

The Chlamydomonas reinhardtii sulfate permease mutant will produce a large amount of 

carbon dioxide that will quickly saturate the process water and begin filling the headspace. 

A piston compressor is used to compress the hydrogen to 300psi. A pressure swing 

adsorption (PSA) system is used to purify the hydrogen gas stream to five nines purity and 

return the CO2 to the CO2 saturation vessel where it will be mixed with a slipstream of water 

that has been removed from raceways low in saturated CO2 during their regeneration phase. 

During the regeneration phase all product gases will be vented. 

Further details of the B-3 pathway are specified in Figure 2-1. The subassemblies of this 

pathway are shown in Figure 4-5. The complete bill of materials and capital costs of this 

production plant are shown in Figure 4-6. The total system cost is $3,562,117. 

Figure 4-5. Production Plant Design for Sulfate Permease (Chlamy)

Photo Bio Reactor Bed

Mixers

Power

Source

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Control System

Subassembly PC Control Room

Instrumentation

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Raceway Wiring Panel

Instrumentation

(x 20 Raceways)

Paddlewheel Motors

(x 20 Raceways)

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Transfer Out

Pump

Recycle

Subassembly

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve

CO2 from

PSA
68 

Figure 4-6. Bill of Materials for B-3 Pathway 

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 165,452 m2 1 m2 $ 0.54 165,452 $ 89,054 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 178,597 m2 1 m2 $ 0.47 178,597 $ 84,382 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 4,000ft2.

Pond Edging 26,566 m 1 m $ 7.00 26,566 $ 185,965 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 38 $ 991,154 Based on California labor rates.

Paddlewheel Mixers 1 each $ 5,000.00 76 $ 380,000 Est. paddlewheel cost. 4 per raceway. 18 raceways.

Inlet Water Valve 0.5 in 1 each $ 43.46 38 $ 1,651 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 38 $ 1,651 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 38 $ 2,555 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 114 $ 912 Price from http://www.ancorp.com/line.aspx?id=819

Gas Capture Subassembly

Compressor 65 kgmol/hr 65 kgmol/hr $ 9,233.00 1 602,582 $ Using H2A Unit cost

PSA $ 80,034.78 1 $ 80,035 Using H2A scaling

Raceway Collection Pipe 76 ft 1 ft $ 2.80 76 $ 213 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

 Main Collection Pipe 2160 ft 1 ft $ 6.18 2160 13,349 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Gas Capture Pipe 350 ft 1 ft $ 6.18 350 2,163 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

 

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 16 164,027 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Nutrient Metering Pumps 1468 gph 1 each $ 2,594 1 2,594 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics. Pumps nutrients into return flow

Mixing Pump 150 gpm 150 gpm $ 10,252 16 164,027 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

CO2 Cylinder 50 lb 50 lb $ 360 1 $ 360 Aluminum 6061-T6 alloy, 1800 psi Service Pressure, 50 lb. Capacity Part #6125 (http://kegman.net/carbon.htm)

CO2 Saturation Vessel 15000 gal 15000 gal $ 30,000 10 300,000 $ Field erected Stainless Steel tank from Perry's chemical Handbook, Plant Design and Economics

Nutrient Tank 15 gal 10 gal $ 30.00 2 $ 60 10 gallon Cylindrical Process Tank, Part #0275-085 (http://www.watertanks.com/products/0275-085.asp)

Raceway Slurry Collection Pipe 76 ft 1 ft $ 2.12 76 $ 161 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 2160 ft 1 ft $ 8.51 2160 18,382 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle-Feed Transfer Pipe 180 ft 1 ft $ 8.51 180 1,532 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle Subassembly

Transfer Out Pump 150 gpm 150 gpm $ 10,252 16 164,027 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Raceway Slurry Collection Pipe 76 ft 1 ft $ 8.51 76 $ 647 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 2160 ft 1 ft $ 8.51 2160 18,382 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Consumables

Initial CO2 41629 lb 50 lb $ 35.00 833 29,155 $ Cost based on average fertilizer use for Aquaculture and average fertilizer costs from the USDA

Initial Nutrients 0.78 lb 1 lb $ 0.20 1 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 3,997,674 gal 1 gal $ 3,997,674 0.0017 $ 6,656 H2A Feedstock Costs

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 8,000 $ comes from price quote of $50/ft2 from http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 3,000.00 $ Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 38 $ 5,548 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 1,500.00 $ Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 4,299.00 $ Information from Emerson Process Management

Level Indicators 1 $ 714.00 38 27,132.00 $ Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 38 13,110.00 $ Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 19 144,400.00 $ Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/analyt_hwhydrogengas.htm)

Air Temperature Meter 1 $ 599.00 1 599.00 $ Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 38 1,444.00 $ Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 38 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 38 16,530.00 $ Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 38 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Instrument Wiring 213180 ft 1 ft $ 0.02 213180 4,135.69 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 264360 ft 1 ft $ 0.02 264360 5,128.58 $ Information from Waytek Inc. at waytekwire.com

Conduit 8820 ft 1 ft $ 0.58 8820 5,113.84 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 3,562,117

4.4 B-4 Immobilized Photobiological System

The B-4 pathway is a dual-bed photobiological reactor system using a sulfur-deprived, 

immobilized, mutant-antenna Chlamydomonas reinhardtii as the H2 production microorganism. H2 production is projected at 3% solar-to-hydrogen conversion efficiency, 

corresponding to a future optimized organism. However, laboratory testing has only yielded 

0.8% conversion efficiency.

The dual-bed system consists of 2 growth reactors and 90 H2 production reactors. Each 

reactor is approximately 40ft wide by 1060ft long. Reactors are large rectangular beds 

without a middle divider because there is no need to direct circular flow. Unlike reactor

systems which use paddlewheels for circulation, these reactors will use a perforated pipeline 

system and central pump to distribute nutrients and provide mixing. The entire system is 

scaled to produce an average of 1,000kgH2/day over the year. All reactors have a liquid 

depth of 10cm53

 

53 10cm depth at 1g/L is not needed for full photon capture. However, the bed is made this deep to provide 

commonality with the other pathways studied, to provide temperature moderation, and to relax construction 

tolerances on the bed itself.

to ensure full photon capture, moderate temperature swings by adding heat 

capacity to the system, and allow for sufficient circulation of nutrients. A thriving algal 
69 

culture is maintained in the growth reactors at an enhanced concentration of >1g/L. Two 

porous, fibrous substrates (2mil spin-molded polypropylene sheets) each 20ft wide and the 

length of the reactor are floated in the growth reactor for 2 days. During this time, the 

newly grown algae attach to the substrate and grow, forming an immobilized biofilm54. 

Commercial fertilizer and CO2 are fed to the growth reactors to ensure micro-organism 

health and growth. After 2 days, the film substrate is transferred to the H2 production 

reactors. 

The substrate then begins to cycle between two days of hydrogen production and two days 

of regenerative growth to keep the culture alive for an extended period of time and to allow 

enough biomass to accumulate as substrate for the respiration needed to keep the system 

anaerobic during hydrogen production. Sulfur deprivation will lead to H2 production, and 

adding sulfur will begin the regeneration phase. During the growth stage, all product gases 

will be vented to the outside. 

The H2 production phase is ~180 days in duration before the film needs replacing. Like in 

the growth beds, the biofilm containing immobilized algae floats to the top of the 10cm 

reactor liquid layer. A custom blend of commercial fertilizer nutrients with a minimal 

amount of sulfur is added to the reactors via a water recirculation system. Sulfur deprivation 

will trigger hydrogen production. As in the growth reactors, a central pump and pipeline 

system circulates nutrients. 

The biofilm substrate from the growth reactors are rolled up and then unrolled into the H2

production reactors using an industrial unwind machine that is attached to the front end of a 

forklift truck. The unwind machine, as quoted by Powell Engineering, costs around $37,000 

and winds the film onto a one foot diameter aluminum core. The Forklift is an $18,571 

Caterpillar C3000 lift. By using the forklift to transport the unwind machine, a single 

device can be used for all biofilm transfers. At the end of the 180-day production cycle, the 

biofilm substrate is discarded. The forklift mounted wind machine is again used to collect 

the substrate. 

The Chlamydomonas reinhardtii micro-organism in sulfur-deprived mode produces a 

nominally pure hydrogen product gas (plus water vapor). A piston compressor is used to 

compress the hydrogen to 300psi. A small pressure swing adsorption (PSA) system is used 

to polish the hydrogen gas stream to five nines purity and return the CO2 to the CO2

saturation vessel where it will be mixed with a slipstream of water that has been removed 

from raceways low in saturated CO2 during their regeneration phase. During the 

regeneration phase all product gases will be vented.

Further details of the B-4 pathway are specified in Figure 2-1. The subassemblies of this 

pathway are shown in Figure 4-7. The complete bill of materials and capital costs of this 

production plant are shown in Figure 4-8. The total system cost is $4,843,599. 

 

54 While a biofilm is “grown” on the substrate, the algae are actually a combination of the free-floating algae 

that clings to the substrate and new algae that blooms during the 2 days of incubation.
70 

Figure 4-7. Production Plant Design for Immobilized Sulfur Deprived

Photo Bio Reactor Beds

(Organism Growth)

Power

Source

Control System

Subassembly

Control Room PC

Instrumentation

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Water

Roll Changing Subassembly

Nutrient / Fertilizer

Tank

Metering

Pumps

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve

L

pH

T

Wiring

3 signal

3 power

Instrumentation

(x 2 Raceways)

Raceway Wiring Panel

Photo Bio Reactor Beds

(H2 Production)

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

8 power

Raceway Wiring Panel

Instrumentation

(x 66 Raceways)

Mixing

Pump

Nutrient

Flow Valve

Roll / Unroll

System

Roll / Unroll

System Forklift for

Transfer

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Roll Changing Subassembly

Control System Subassembly

Sensor / Alarms

Indicators

Organism Feed Subassembly

Transfer Out

Pump

Recycle

Subassembly Slip

Stream
71 

Figure 4-8. Bill of Materials for B-4 Pathway 

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 381,075 m2 1 m2 $ 0.54 381,075 $ 205,113 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 411,355 m2 1 m2 $ 0.47 411,355 $ 194,354 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 4,000ft2.

Pond Edging 61,258 m 1 m $ 7.00 61,258 $ 428,807 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 92 $ 2,399,636 Based on California labor rates.

Inlet Water Valve 0.5 in 1 each $ 43.46 90 $ 3,911 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 90 $ 3,911 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 90 $ 6,051 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 270 $ 2,160 Price from http://www.ancorp.com/line.aspx?id=819

Gas Capture Subassembly

Compressor 56 kgmol/hr 56 kgmol/hr $ 9,233.00 1 521,146 $ Using H2A Unit cost

PSA $ 60,970.48 1 $ 60,970 Using H2A scaling

Raceway Collection Pipe 180 ft 1 ft $ 1.00 180 $ 180 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Main Collection Pipe 5280 ft 1 ft $ 6.18 5280 32,630 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

 Gas Capture Pipe 460 ft 1 ft $ 6.18 460 $ 2,843 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

 

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 1 10,252 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Nutrient Metering Pumps 1468 gph 1 each $ 2,594 1 $ 2,594 Est. Based on Perry's Chemical Handbook, Plant Design and Economics. Pumps nutrients into return flow

Mixing Pump 150 gpm 150 gpm $ 10,252 1 10,252 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

CO2 Cylinder 50 lb 50 lb $ 360 1 $ 360 Aluminum 6061-T6 alloy, 1800 psi Service Pressure, 50 lb. Capacity Part #6125 (http://kegman.net/carbon.htm)

CO2 Saturation Vessel 15000 gal 15000 gal $ 30,000 1 30,000 $ Field erected Stainless Steel tank from Perry's chemical Handbook, Plant Design and Economics

Nutrient Tank 35 gal 10 gal $ 30.00 4 $ 120 10 gallon Cylindrical Process Tank, Part #0275-085 (http://www.watertanks.com/products/0275-085.asp)

Nutrient Piping 90000 ft 1 ft $ 0.58 90000 52,200 $ 

Raceway Slurry Collection Pipe 180 ft 1 ft $ 0.52 180 $ 94 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 5280 ft 1 ft $ 2.12 5280 11,194 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle-Feed Transfer Pipe 180 ft 1 ft $ 2.80 180 $ 504 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Roll Changing Subassembly

Roll-Unroll System 20 ft 1 Roller $37,000 1 37,000 $ Information from conversations with Powell Engineering. $29,500 for base, $4,000 for motor, and $3,500 for Core 

Polypropylene Film 352070 m2 1 m2 $ 0.12 352070 42,248 $ 

Forklift 1 Forklift $ 18,571.00 1 18,571 $ Price Quote from Atlantic Lift Truck Inc. 

Recycle Subassembly

Transfer Out Pump 150 gpm 150 gpm $ 10,252 1 10,252 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Main Slurry Feed Pipe 5280 ft 1 ft $ 2.12 5280 11,194 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Raceway Slurry Feed Pipe 180 ft 1 ft $ 0.52 180 $ 94 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Consumables

Initial CO2 206180 lb 50 lb $ 35.00 4124 144,340 $ Cost based on average fertilizer use for Aquaculture and average fertilizer costs from the USDA

Initial Nutrients 1.81 lb 1 lb $ 0.20 2 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 9,324,121 gal 1 gal $ 9,324,121 0.0017 $ 15,525 H2A Feedstock Costs

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 $ 8,000 comes from price quote of $50/ft2 from http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 $ 3,000 Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 90 $ 13,140 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 $ 1,500 Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 $ 4,299 Information from Emerson Process Management

Level Indicators 1 $ 714.00 90 $ 64,260 Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 90 $ 31,050 Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 45 $ 342,000 Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/analyt_hwhydrogengas.htm)

Air Temperature Meter 1 $ 599.00 1 $ 599 Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 90 $ 3,420 Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 90 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 90 $ 39,150 Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 90 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 $ 5,500 Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 $ 5,500 Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 $ 5,500 Information from Emerson Process Management

Instrument Wiring 1065720 ft 1 ft $ 1065720 20,675 0.02 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 1327500 ft 1 ft $ 1327500 25,754 0.02 $ Information from Waytek Inc. at waytekwire.com

Conduit 20260 ft 1 ft $ 0.58 20260 11,747 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 4,843,599

4.5 B-5 Pathway Production Plant

The B-5 pathway is a Chemostat II photobiological reactor system using a mutant-antenna 

purple-non-Sulfur (PNS) proteobacteria as the H2 production micro-organism. As a 

Chemostat II, the initial bacterial culture is grown at the normal cell growth rate until final 

colony concentration is reached, and then the process is switched over to steady-state mode 

where simultaneous growth and hydrogen production occur in the reactor beds. This is 

achieved by manipulating the nutrients. H2 production is projected at 3.5% solar-tohydrogen efficiency and corresponds to a future optimized organism. Current PNS systems 

have been demonstrated at only 2.5% conversion efficiency in the laboratory. Sufficient 

cellular activity to keep the micro-organism healthy is assumed to occur during H2

production with new cell growth taking place at a reduced rate compared to wild-type 

organisms (20% per day vs. ~80% per day). Since the system functions as a Chemostat II 

with 20% daily organism renewal, an effectively infinite H2 production phase duration is 

achieved. 

Fifty-four raceways of approximately 40ft width by 1090ft length are used to produce an 

average of 1,000kgH2/day throughout the year. The raceways have a liquid depth of 10cm 

to correspond to full photon capture at a cell concentration of 0.2g/L. Paddlewheels 

circulate the water/PNS slurry. A slip stream of water/micro-organism slurry is drawn off 
72 

continuously to maintain the 0.2g/L micro-organism concentration. Micro-organisms are 

removed from the recycle stream by a rotary drum filter. The micro-organism mass 

removed is carefully matched to the growth rate to maintain a constant organism mass 

within the system. The removed micro-organisms are taken to a land-fill, used in a 

subsequent fermentor, or otherwise disposed. Nutrients in the form of acetate are mixed 

with the returning water to provide required nourishment to the micro-organisms and as a 

substrate in the hydrogen production reaction. 

The PNS micro-organisms produce a net product gas of 95% H2 and 5% CO2 (plus water 

vapor). A piston compressor is used to compress the gas mixture to 300psi prior to 

separation. A pressure swing adsorption (PSA) system is used to purify the hydrogen gas 

steam.

Further details of the B-5 pathway are specified in Figure 2-1. The subassemblies of this 

pathway are shown in Figure 4-9. The complete bill of materials and capital costs of this 

production plant are shown in Figure 4-10. The total system cost is $4,034,192. 

Figure 4-9. Production Plant Design for Purple Non-Sulfur

Photo Bio Reactor Bed

Mixers

Power

Source

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Control System

Subassembly

Control Room PC

Instrumentation

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Raceway Wiring Panel

Instrumentation

(x 40 Raceways)

Paddlewheel Motors

(x 40 Raceways)

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Transfer Out

Pump

Wet Spent

Organisms

Makeup

Water

Algal Separation

Recycle

Subassembly

CO2

Cylinder

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve
73 

Figure 4-10. Bill of Materials for B-5 Pathway 

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 235116 m2 1 m2 $ 0.54 235,116 $ 126,551 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 253795 m2 1 m2 $ 0.47 253,795 $ 119,912 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 4,000ft2.

Pond Edging 37752 m 1 m $ 7.00 37,752 $ 264,266 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 54 $ 1,408,482 Based on California labor rates.

Paddlewheel Mixers 1 each $ 5,000.00 108 $ 540,000 Est. paddlewheel cost. 4 per raceway. 18 raceways.

Inlet Water Valve 0.5 in 1 each $ 43.46 54 $ 2,347 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 54 $ 2,347 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 54 $ 3,630 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 162 $ 1,296 Price from http://www.ancorp.com/line.aspx?id=819

Gas Capture Subassembly

Compressor 45 kgmol/hr 45 kgmol/hr $ 9,233.00 1 414,241 $ Using H2A Unit cost

PSA $ 30,541.73 1 30,542 $ Using H2A scaling

Raceway Collection Pipe 108 ft 1 ft $ 1.00 108 $ 108 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Main Collection Pipe 3120 ft 1 ft $ 6.18 3120 19,282 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

) Gas Capture Pipe 50 ft 1 ft $ 6.18 50 $ 309 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

 

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 3 30,755 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Nutrient Metering Pumps 1468 gph 1 each $ 2,594 1 2,594 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics. Pumps nutrients into return flow

Mixing Pump 150 gpm 150 gpm $ 10,252 3 30,755 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Nutrient Tank 21 gal 10 gal $ 30.00 3 $ 90 10 gallon Cylindrical Process Tank, Part #0275-085 (http://www.watertanks.com/products/0275-085.asp)

Raceway Slurry Collection Pipe 108 ft 1 ft $ 1.00 108 $ 108 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 3120 ft 1 ft $ 4.31 3120 13,447 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle-Feed Transfer Pipe 180 ft 1 ft $ 6.18 180 1,112 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle Subassembly

Rotary Drum filter 25741 gph 3750 gph $ 87,000.00 7 609,000 $ Information from Dana Kent at Advanced Aquaculture Inc. on Hydrotech Drum Filter

Transfer Out Pump 150 gpm 150 gpm $ 10,252 3 30,755 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics/Peters and Timmerhaus

Main Slurry Feed Pipe 3120 ft 1 ft $ 4.31 3120 13,447 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Raceway Slurry Feed Pipe 108 ft 1 ft $ 1.00 108 $ 108 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Consumables

Initial Nutrients 1.11 lb 1 lb $ 0.20 1 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 5,720,312 gal 1 gal $ 5,720,312 0.0017 $ 9,524 H2A Feedstock Costs

Initial Acetate 7,793 kg 1 kg $ 0.5950 7,793 $ 4,637 Chemical Journal of Korea. Assumes consumption to be four times more to account for full growth conditions.

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 8,000 $ comes from price quote of $50/ft2 from http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 3,000.00 $ Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 54 $ 7,884 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 1,500.00 $ Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 4,299.00 $ Information from Emerson Process Management

Level Indicators 1 $ 714.00 54 38,556.00 $ Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 54 18,630.00 $ Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 27 205,200.00 $ Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/analyt_hwhydrogengas.htm)

Air Temperature Meter 1 $ 599.00 1 599.00 $ Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 54 2,052.00 $ Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 54 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 54 23,490.00 $ Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 54 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Instrument Wiring 406380 ft 1 ft $ 0.02 406380 7,883.77 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 505080 ft 1 ft $ 0.02 505080 9,798.55 $ Information from Waytek Inc. at waytekwire.com

Conduit 12340 ft 1 ft $ 0.58 12340 7,154.73 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 4,034,192

5.

Thus far in the report we have discussed the basic science that applies to photobiological 

systems as well as the biological parameters of micro-organisms and the engineering design 

conditions which were used in the analysis. This section of the document is dedicated to 

financial and engineering assumptions that have been made in order to evaluate the 

economics of each of the bio-hydrogen solutions under review. The assumptions have been 

divided between those necessary to develop capital costs and those essential to determining 

levelized hydrogen costs. That is followed by computation of the levelized hydrogen costs 

and a discussion of the results and how they may be improved. 

Capital Cost Assumptions and Calculations

5.1 Capital Cost Assumptions

A bill of materials for each pathway outlining the capital costs associated with the 

construction and components of that photobiological hydrogen production facility is 

provided in the previous section. That section characterized the basic operation and 

configuration assumptions but many other assumptions relating to the capital cost were not 

defined. This section outlines and defines the other underlying assumptions used in 

determining the capital costs associated with the individual cost components. 
74 

5.1.1 Material Margins

When calculating the size of the Low Density Polyethylene (LDPE) film needed to cover 

the top of the raceways, it was necessary to add an extra margin to account for ease of 

assembly and the potential for misalignment as the film is unrolled. We assumed an 

additional one foot would be more than sufficient to cover the pond edging and leave extra 

material for sealing against the edging. In addition, we calculate that an additional six inches 

on either side and on either end will be necessary to give room for the accumulation of gas

at times of day when production of gas exceeds removal. Excess gas accumulation is

covered in Section 5.2.3. 

5.1.2 Bed Dimensions

In choosing the dimensions of the raceways, polyethylene manufacturers were consulted to 

determine the manufacturing constraints of polyethylene film production. Since the LDPE is 

intended as an impermeable hydrogen gas barrier, it is desirable to avoid seams between 

sheets as those are at higher risk for leakage. While sealing on the edges of the raceways is 

unavoidable, the difficulty of sealing against hydrogen loss led us to size the raceways based 

on the largest single sheet of LDPE that could be produced. Conversations with Berry 

Plastics indicated that the maximum width for a roll of LDPE film is 56 feet. However, we 

recognized the constraints of transporting such a large roll, and assumed a maximum width 

that could fit onto a truck to be roughly 50 feet. To allow for the overage mentioned in the 

previous section, we assume a roll width of 43 feet, leaving a 40 foot wide raceway. Berry 

Plastics also indicated that roll lengths of over 1000 feet are feasible. To ensure 

impermeability to hydrogen, Berry Plastics recommended 6mm thick film. We desired an 

even number of raceways to make material cost calculations and system design easier. 

Therefore, to determine the length of raceways for each system, we chose a round number of 

feet near 1000 that, with a 40 foot active width, would cover the necessary production area 

with an even number of raceways. For all systems, raceway length is between 1060 and 

1090ft. In addition to these physical dimension parameters, the LDPE also impacts bed 

dimension because it is a barrier to the full solar insolation reaching the algae. Data from 

manufacturer indicates that the transmission of sunlight is 90% across the film. That has 

been taken into consideration in sizing the raceways. 

5.1.3 Gas Capture Sizing

Due to yearly variation in insolation, it is necessary to size the elements of the gas capture 

subassembly for the projected peak day of production, June 21. Sizing for the average 

would lead to extreme H2 accumulation of during the days where insolation is higher than 

the average (March 21 to September 21). Thus, the gas capture subassembly will be 

operating below rated capacity for most of the year. 

5.1.4 Pipe Sizing

Pipe sizing was achieved through the use of the continuity equation = ρAv. For gas 

piping, we assumed a maximum gas velocity of 100 ft/second to limit pressure loss due to 

pipe flow. For water piping, we assumed a maximum velocity of 15 ft/second. The exact 

physical layout of the piping systems is needed to accurately assess the velocity and 

pressure drop relationship. Since a detailed piping layout is beyond the scope of this 

project, the specific velocity cannot be known. Consequently, the number computed is used 

as a surrogate for more detailed analysis and we judge it to be adequate given the low costs 
75 

associated with small changes in PVC pipe diameter. We rounded our pipe diameter 

calculated to the nearest nominal size for cost analysis purpose. 

There were three different pipe sizes calculated for the various stages of liquid removal and 

recycling. The pipes that transfer the water out to and back from the individual raceways 

are sized for individual raceway requirements. Each of these individual raceway pipes 

comes from or combines with a collection manifold sized for half of the overall volume, 

since the raceways are placed in two even rows with the recycle and algae feed 

subassemblies in the middle. Finally, one pipe combines flow from the two collection 

manifolds and feeds it to the recycle subassembly and out of the algae feed subassembly. 

This manifold must be sized to contain the entire flow volume. The pipe sizes for each 

system are given in Figure 5-1. Similarly, gas piping was sized both for individual outlets 

on each raceway, and then the main collection lines that lead to the Gas Capture 

Subassembly. 

Figure 5-1. Piping sizes for B-1 through B-5 

Algal O2-

tolerant 

Hydrogenase 

Cyanobacterium

O2-tolerant 

Hydrogenase 

Algal

Sulfate

Permease

Immobilized 

Algal, Sulfur 

deprived

PNS 

bacteria

B-1 B-2 B-3 B-4 B-5

Gas Piping

Raceway 

Collection Pipe 1.5” 1.5” 1” 1” 1” 

Main Collection 

Pipe 4” 4” 4” 3.5” 3.5” 

Gas Capture 

Collection Pipe 5.5” 5.5” 5.5” 5.0” 4.5” 

Slurry Piping

Raceway Slurry 

Collection 1” 1” 2” 0.5” 1” 

Main Slurry 

Collection 2” 2” 7” 2” 3” 

Recycle-Feed 

Transfer Pipe 3” 3” 10” 2.5” 4.5” 

Main Slurry 

Feed Pipe 2” 2” 7” 2” 3” 

Raceway Slurry 

Feed Pipe 1” 1” 2” 0.5” 1” 

5.1.5 Equipment Duty Cycles

Given that the biological production of hydrogen gas can only occur during sunlight, it was 

originally assumed that the operation cycles of the gas capture subassembly equipment 

would match the production cycle. However this would necessitate sizing the compressor 

and the PSA for the maximum gas production rate. Considering the substantial capital and 

operation costs associated with the compressor and the PSA, it is more cost effective to size 

the compressor and the PSA for average daily gas production and have it run 24 hours a day 

so the last of the product gases are removed from the bed as a new production day begins. 
76 

The organisms require nutrients even though they may not be producing hydrogen. Because 

of this the replenishment supply and slip stream removal must be continuous. Thus the 

pumps and drum filters are also expected to operate 24 hours a day and sized as such. 

Paddlewheel operation is not continuous because the mixing of the micro-organisms is only 

required during sunlight hours. 

5.1.6 Operational Cycle

While ideally the hydrogen production bioreactor would operate all days of the year, it will 

be necessary to shut down the reactor at times for maintenance or to flush out and restart 

individual raceway colonies in case of culture contamination. With these considerations in 

mind, we have assumed a plant operation time of 90%. 

In a given day, the production cycle will be entirely determined by the duration of sunlight, 

which varies throughout the year. To simplify the equipment sizing procedure we have 

assumed that on average, there will be 12 hours of sunlight per day throughout the year. 

5.2 Calculations Explained

Many of the calculations used for the sizing of elements in the raceway are explained in the 

following sections. They are based on assumptions stated in the previous section or 

elsewhere in this report. 

5.2.1 Quantity of Raceways 

As mentioned in Section 1.1.2 the dimensions of the raceways are determined largely by the 

constraints of the LDPE film production and practical considerations such as ease of truck 

transport. We slightly tweaked the length in order to obtain an even number of raceways. 

However, with these numbers relatively constrained, the determining factor for quantity of 

raceways is the area of bed required to produce 1,000kg of usable hydrogen gas per day. 

According to PSA modeling and conversations with PSA manufacturer UOP, PSA systems 

operating on H2/O2 gas mixtures can achieve approximately 90% hydrogen recovery. As 

mentioned earlier, the LDPE film is only 90% transparent as well. Thus we must oversize 

the reactor bed in order to account for both this photon loss and PSA hydrogen losses and 

achieve the target net 1,000kgH2/day. Based on the efficiency assumptions of the particular 

system involved, the average insolation data referenced earlier, and the assumed conversion 

efficiency of photons to hydrogen also mentioned earlier, we integrated the hourly H2

production to get a daily gH2/m2 rate. This, combined with our desired H2 production level, 

allowed us to calculate the total production area required. Knowing the constraints of our 

raceway size mentioned earlier, we determined the quantity of raceways by dividing the 

total area needed by the size of an individual raceway. 

5.2.2 Water Slip Stream Volume or Rates

Although organisms in hydrogen production mode generally grow at a reduced rate, there is 

some continual growth occurring. In order to maintain a constant concentration of 

organisms within the bed it is necessary to draw off a specific volume of water/organism 

solution to filter out the excess organisms. The specific slip stream volume required is 

based on the average doubling rate of our organisms. Using hourly solar intensity data and 

the corresponding organism growth rate, an iterative process determines that a removal rate 
77 

of 0.45%/hour of the total bed aqueous volume is needed to ensure that concentration is 

constant throughout the day. Each system has a different total reactor volume depending on 

the assumed efficiency of hydrogen production in the system. Making use of 0.45% of the 

overall volume of the system as the hourly removal mass fraction, the hourly slip stream 

volume is computed. This volume value is employed to size pumps, pipes, and saturation 

vessels. In the B-3 and B-4 system a slip stream of water will be pumped to the CO2

saturation vessel for CO2 addition necessary for the growth and regeneration phase. The 

volume of water drawn off was calculated by determining how much water is needed each 

day of the regeneration phase to fully absorb all the CO2 that was bubbled into the 

headspace during the hydrogen production phase based on the CO2 saturation rate of 1.4g/L. 

5.2.3 Product gas storage under Film

As mentioned in Section 1.1.4, to keep compressor and PSA size down, the rate of gas 

drawn off of the raceways will not necessarily match the rate of gas production. As a 

consequence, excess gas will accumulate underneath the LDPE film. This accumulation will 

necessitate a surplus headspace volume above the liquid level of the bed and below the 

LDPE film. The LDPE film above the bed will rise as the gas accumulates during the day 

and fall as the gas is drawn out during the night. An excess of film material is needed to 

accommodate this film motion and avoid stretching of the film. In order to calculate the 

amount of extra film necessary, we used data from the NREL SOLPOS model combined 

with data from the NASA Atmospheric Data Center to calculate the amount of hydrogen gas 

being produced at 15 minute intervals during the month of June where insolation and 

corresponding hydrogen production will be at its peak55, 56. Based on the molar ratios of 

each system, we calculated the amount of total gas produced for each 15 minute interval. 

Knowing the density and the molar ratios of the gas we calculated the volume. Using the 

volume of gas and the area of the raceway, we computed the total height required of the film 

to achieve that volume. Lastly, by subtracting a specific rate of gas every 15 minutes, we 

determined the minimum rate of gas removal that will leave no gas accumulated by the 

following dawn. According to our calculations, approximately 6cm of headspace is required 

to account for gas accumulation. The resulting gas removal rate is then used in PSA sizing.

5.2.4 Electricity Consumption

The power consumption of the system primarily comes from the Gas Capture Subassembly, 

with slight contributions elsewhere. The items consuming power are the compressor, PSA, 

pumps, nutrient metering pumps, rotary drum filters, paddlewheels and the control room. 

PSA power is minimal as it is only needed for air compression to actuate pneumatic valves. 

The power consumption of the pumps, nutrient metering pumps, rotary drum filters and 

paddlewheels all come from quotes we have received concerning these items and have been 

documented previously in this report. The gas compressor is assumed to be a 2-stage piston 

compressor (N=2) with interstage cooling. Its power was calculated by assuming isentropic 

compression from 20oC ambient temperature with efficiency, η, of 75%. Overall pressure 

ratio was 20.4 for an outlet pressure of 300psi. 

 

55 NREL MIDS SOLPOS (Solar Position) model.

56 NASA Atmospheric Data Center, Langley ASDC User Services, Surface meteorology and Solar Energy 

(SSE) data base (release 6.0).
78 

Lastly, the power consumption of the control room was calculated by taking 5% of the 

overall power consumption of the plant. Given the relatively low power consumption of the 

control room and the ambiguity of all the different power consuming devices it might 

contain, this fairly conservative estimate is adequate. Figure 5-2 lists the electricity usage 

anticipated for each of the components previously described and provides a total 

consumption value to be used in further analysis. The previously mentioned duty cycles and 

operational cycle have been taken into account in these computations.

Figure 5-2. Electricity Consumption per Pathway57

B1 B2 B3 B4 B5

Compressor kW 123 123 78 78 83

PSA kW 1 1 1 1 1

Pumps (3) kW 18 18 96.0 9 27

Nutrient Pump kW 1.5 1.5 1.5 1.5 1.5

Control Room kW 9.8 9.8 13.5 4.5 12.7

Rotary Drum Filters kW 3.9 3.9 0.0 0 9.1

Paddlewheels kW 49.2 49.2 93.5 0 132.8

Total kW 206.4 206.4 283.7 94.2 267.0

Total Usage kWh/yr 1,433,023 1,433,023 2,041,916 742,765 1,581,043

Usage Rate kWh/kg H2 4.36 4.36 6.22 2.26 4.81

Cost $/yr $70,957 $70,957 $101,106 $36,778 $78,286

Cost $/kg H2 $0.20 $0.20 $0.28 $0.10 $0.22

5.2.5 Excavation of Land for Reactor Bed Placement

Because of the considerable size of the raceways, detailed attention was given to calculating 

an accurate cost for the construction of the reactor ponds. Consultation with Mark 

Dormsteader from Metro Earth Works, a company that focuses solely on earth moving 

projects, suggested that our project would require a loader, a dozer, and a roller. He said 

that medium sized equipment would be adequate for such a shallow pool depth. The project 

would also require a foreman and two laborers on foot. No dump truck for dirt removal will 

be necessary since all dirt excavated will be used in construction of the side and central

berms. Assuming a standard 8 hour work day, it was estimated to take 5 days to construct 

one raceway. Since multiple teams of workers would be working simultaneously to build 

all the raceways necessary for a plant, the time per raceway constructed would be 

significantly reduced, however, the overall cost per raceway would remain the same, since 

increased workforce brings increased cost. 

To calculate the wages of the workers, we used the Department of Labor’s Davis-Bacon 

Wage Determinations for 2008, as referenced in the US Department of Agriculture’s “Cost 

Estimating Guide for Road Construction”, which provide a state by state breakdown of 

 

57 Equipment consumptions rates are based on producing of 1,111kgH2/day. Yearly Consumption is based on 

all equipment operating 360days/yr, 24hrs/day, with the exception of paddlewheels, which operate only 

12hrs/day. Consumption Rate/Hydrogen Product is based on final product of 1,000kgH2/day over previously 

stated 360 days. 
79 

average wages for various jobs58

Figure 5-3

. We also used equipment rental costs referenced by this 

report taken from The Blue Book of Building and Construction. The cost for equipment

rental and the wages vary widely, causing substantial variability in the construction price 

depending on location. shows some examples of wages and rental costs in the 

American Southwest. We chose the numbers corresponding to Californian wages and rental 

rates. Given the fact that the organisms involved in biohydrogen production are sensitive to 

temperature, we decided that a location with more moderate temperatures, yet still high 

levels of sunlight would be ideal. Thus, southern California, which has more moderate 

temperatures due to its proximity to the ocean, would be more ideal than desert areas in 

Arizona or Nevada. Although not considered in our analysis, the proximity to the ocean 

could also provide an added benefit: an endless supply of water for the system59. Figure 5-4

shows the cost estimation for the B-1 system, which requires 20 raceways. Were the plant 

built in some other state, such as New Mexico, there would be up to a 50% reduction in 

construction costs. 

Figure 5-3. Davis-Bacon Wage Determinations and Blue Book Rental Costs 

 

 

58 Moll, Jeff, Marjorie Apodaca, Ken Goddard, Jon Stites, Andrea Glover. Cost Estimating Guide for Road 

Construction. US Forest Service, USDA, Washington DC. April 2008.

59 While some organism may flourish in saltwater, we assume fresh water for all of the biohydrogen systems 

under analysis. Consequently, a desalination plant would be necessary to achieve unlimited water from the 

ocean.
80 

Figure 5-4. Excavation Cost Estimate for B-1 System using California Costs 

Equipment Cost- Loader 93.68 $/hr

Equipment Cost- Dozer 94.25 $/hr

Equipment Cost- Roller 74.3 $/hr

Operator Cost- Loader 71.38 $/hr

Operator Cost- Dozer 71.38 $/hr

Operator Cost- Roller 68.31 $/hr

Laborer Cost 49.97 $/hr

Foreman Cost 78.84 $/hr

Number of Laborers 2

Number of Operators 3

Total Cost/ day $ 5,217

Total Cost/ pool $ 26,083

Total Cost/ system $ 521,664

Excavation Cost Estimation for California

The excavation costs calculated concern only the construction of the raceways and the 

berms. The installation costs of the other components in the system are computed using the 

H2A methodology. 

6.

Thus far, this report has discussed the capital cost and expenditures associated with building 

a photobiological-to-hydrogen plant. While this is critical and valuable information, the 

build decision is ultimately based on the financial and economic benefits of these capital 

expenditures. The investment is quite large and cannot be properly evaluated without some 

knowledge of the expected return on that investment. In order to evaluate the return, DTI 

performed a discounted cash flow (DCF) analysis using the H2A Production Model, 

Version 2.0. 

Levelized Costs Assumptions & Calculations (H2A)

The H2A model provides a structured format for a user to enter in parameters which impact 

cash inflows and outflows associated with the construction and operation of the plant. 

There are H2A Default values for several of the parameters which can be selected as part of 

the analysis. Additionally, there are plant specific parameters which the analyst must enter. 
81 

Once all parameters have been entered the H2A model computes the levelized cost of 

hydrogen in $/kgH2.

60

6.1 Standard H2A parameters

In order to develop levelized costs, several parameters must be defined. Because this 

analysis focuses on a plant which is still in its conceptual stage, many of the values for these 

parameters must be assumed. The assumptions are documented in this section. These are 

meant to represent a baseline system and analysis. Later in the discussion portion of this 

report, the assumed values can be altered for sensitivity analyses.

Standard H2A financial values and assumptions shown in Figure 6-1 apply to all 

biohydrogen pathways. This list does not encompass all parameters which must be defined 

in order to run the analysis; just those were there is an H2A Default value which has been 

accepted for this analysis. The remaining parameters are defined later in this section. Some 

of those parameters are common to all biohydrogen pathways and others are pathway 

specific. No dispensing parameters are listed here because the biohydrogen plants are 

central type plants and thus dispensing is not factored into the analysis. 

 

60 For further description of the H2A Model reference, D. Steward, T. Ramsden, and J. Zuboy. H2A 

Production Model, Version 2 User Guide. NREL/TP-560-43983. Golden, CO. September 2008.
82 

Figure 6-1. H2A Default Values and Assumptions used for all Biohydrogen Pathways61

Parameter Assumptions

Analysis Period 20 years 

Burdened Labor Rate for Staff $50/hour 

CO2 Capture Credit Not included in base cases (default 

value = 0) 

CO2 Production Taxes Not included in base cases (default 

value = 0) 

Construction Period and Cash Flow 1 year

Co-produced and Cogenerated Electricity 

Price 

$30/MWh 

Decommissioning 10% of initial capital

Depreciation Type and Schedule for Initial 

Depreciable Capital Cost 

MACRS: 20 years for central model 

Facility Life 20 years 

G&A Rate 20% of the staff labor costs 

Hydrogen Pressure at Central Gate 300 psig

Hydrogen Purity62 98% minimum; CO < 10 ppm, sulfur < 

10 ppm 

Income Taxes 35% federal; 6% state; 38.9% effective 

Inflation Rate 1.9%, but with resultant price of 

hydrogen in reference year constant 

dollars 

Installation Cost Factor 1.3

Land Cost $5,000/acre purchased 

Licensing, Permits and fees $1000 

O2 Credit Not included in base cases 

Production Maintenance & Repairs 1⁄2% of direct capital cost 

Property Taxes and Business Insurance 2%/year of the total initial capital cost 

Reference Financial Structure 100% equity with 10% IRR; includes 

levelized hydrogen price plot for 0%–

25% IRR; model allows debt financing 

Sales Tax Not included on basis that facilities and 

related purchases are wholesale and 

through a general contractor entity 

Salvage Value 10% of initial capital

Working Capital Rate 15% of the annual change in total 

operating costs 

 

61 D. Steward, T. Ramsden, and J. Zuboy. H2A Production Model, Version 2 User Guide. Appendix 3: Default 

Values and Assumptions. NREL/TP-560-43983. Golden, CO. September 2008. p. 60. 

62 Purity levels are driven by H2 vehicle consumption requirements.
83 

6.2 Pathway Common Parameters

In addition to the financial parameters defined by H2A in the previous section there are 

other project inputs which must be quantified in order to carry out the DCF analysis. All 

inputs can be found on the following worksheets in the H2A model; 

• Input_Sheet_Template 

• ReplacementCosts

• CapitalCosts

Many of these parameters are specific to the location, operation, and type of plant. In the 

case of our biohydrogen pathways there are some parameters that are the same for all 

pathways and some that vary by pathway. The parameters in Figure 6-2 are common to all 

pathways. 

Figure 6-2. Parameters Common to all Pathways 

Parameter Assumed Value Worksheet

6.2.1 Operating Capacity Factor

The operating capacity factor can be found on the Input_Sheet_Template worksheet of the 

model. This analysis assumes that the plant and dispensing station have a 90% operating 

capacity. This capacity factor takes into considering things such as planned maintenance 

outages, forced outages, etc. Thus, if the plant is capable of producing 1,111kgH2/day, only 

1,000kgH2 will be dispensed and all economic benefit analysis is based on the amount 

dispensed or sold. 

6.2.2 Reference Year Dollars

The reference year dollars parameter is the year dollars in which the cost of hydrogen is 

reported. The H2A standard is to report out hydrogen costs in 2005 dollars. The model 

expects capital costs to be entered in 2005 dollars. In this analysis Reference Year 2005 

was selected.

6.2.3 Site Preparation

The site preparation parameter can be found on the Input_Sheet_Template worksheet of the 

model. In central plants, H2A defaults this value to 1% of direct costs. This analysis uses 

the same default value; however, the cost basis is slightly altered. Our direct costs include 

excavation for raceways. We excluded that from the cost basis in this calculation since it is 

Operating Capacity Factor 90% Input_Sheet_Template

Reference Year Dollars 2005 Input_Sheet_Template

Site Preparation 1% of direct costs minus 

raceway excavation 

costs 

Input_Sheet_Template

Engineering & design 7% of direct costs Input_Sheet_Template

Process Contingency 20% of direct costs Input_Sheet_Template

Project Contingency $0 Input_Sheet_Template

Up-Front Permitting Costs 0.5% of direct costs Input_Sheet_Template

Production Maintenance & 

Repairs

0.5% of direct costs Input_Sheet_Template
84 

not a piece of equipment, but rather a separately computed type of site preparation. 

Excavation is not all inclusive as site preparation is still required for things such as 

driveways, control building and algae feed subassembly. 

 

6.2.4 Engineering & design

The engineering & design parameter can be found on the Input_Sheet_Template worksheet 

of the model. In central plants, H2A defaults this value to 13% of direct costs. This 

analysis uses 7% of direct costs due to the modularity of the design. 

6.2.5 Process Contingency

The process contingency parameter can be found on the Input_Sheet_Template worksheet 

of the model. In central plants, H2A defaults this value to 15% of direct costs. This 

analysis uses 20% of direct costs due to uncertainties in the system configuration. 

6.2.6 Project Contingency

The process contingency parameter can be found on the Input_Sheet_Template worksheet 

of the model. In our analysis we have chosen to include all contingency factors in the 

process contingency parameter, thus the project contingency is set to $0. 

6.2.7 Up-Front Permitting Costs

The up-front permitting cost parameter can be found on the Input_Sheet_Template 

worksheet of the model. The H2A default for this parameter is 9% of direct costs. This 

analysis uses 0.5% of direct costs due to the modularity of the design and its environmental 

benefits.

6.3 Pathway Specific Parameters

The last types of parameters we identify are those which are specific to each pathway 

analyzed. Figure 6-3 lists these parameters and rules of thumb applied in computing their 

values. These are pathway specific because they are associated with feedstock, process 

design and plant design. 

Figure 6-3. Pathway Specific Parameters

Parameter Rule Applied Worksheet

Land Required 30% greater than reactor 

bed area

Input_Sheet_Template

Production facility plant staff 1 worker per shift per 

100 raceways + 0.5

workers for winders

Input_Sheet_Template

Utility Usage Electricity and water 

costs use H2A pricing

Input_Sheet_Template

Feedstock Usage As required by organism Input_Sheet_Template

Specified Yearly Replacement 

Costs

LDPE every 5 yrs ReplacementCosts
85 

6.3.1 Land Required

The land required parameter can be found on the Input_Sheet_Template worksheet of the 

model. In developing the plant design, a land area for the reactor bed was computed. The 

analysis assumes that the total land required for this design is 30% greater than the reactor 

bed. The 30% factor is meant to encompass area requirements for pump skids, compressors, 

gas and algae separators and a small control room. The total land requirement for each 

pathway is shown in Figure 6-4. 

Figure 6-4. Land Required for each Pathway

Algal O2-

tolerant 

Hydrogenase 

(1)

Cyanobacterium 

O2-tolerant 

Hydrogenase (2) 

Algal 

Sulfate 

Permease

Immobilize

d Algal, 

Sulfur 

deprived

PNS 

bacteria

B-1 B-2 B-3 B-4 B-5

Bed Area

(m2) 

80,968 80,968 151,035 352,070 216,228

% increase 

for 

auxiliaries

30% 30% 30% 30% 30% 

Total Land 

Required 

(m2)

105,259 105,259 196,345 457,691 281,096 

6.3.2 Production facility plant staff

The production facility plant staff parameter can be found on the Input_Sheet_Template 

worksheet of the model. This represents the number of full-time employees required to 

operate the plant. There is no H2A default value for this. Our analysis for each of the 

pathways is shown in Figure 6-5. From our analysis we suggest 3 employees are needed for 

each 1,000kgH2 module of the B-1, B-2, B-3, and B-5 system design. This is based on the 

assumptions that each worker can operate 5 of these modules simultaneously but the plant 

operates 24 hrs/day so three shifts are needed. In the case of the B-4 pathway, an additional 

employee is required for the manual transfer of algae films between ponds. However, the 

films only need replacement once every six months, limiting the overall labor needed.. We 

have assumed a wind and unwind speed of 0.5 ft/second for the polypropylene films. Two 

old films per raceway will need to be wound and unwound along with new films that will be 

wound and unwound for replacement. Additionally, we have assumed it will take 15 

minutes to reposition and prepare the winder for each wind/unwind cycle. We calculated 

that it will take about eight hours, or one full FTE for a worker to complete the task of 

exchanging the mats in one raceway. Thus, the total labor need is the total number of 

raceways (90) divided by the number of work days in 6 months (120), giving us ~0.75 FTE 

in addition to the daily raceway operators. 
86 

Figure 6-5. Plant Staff Requirements for 1 tonne H2/day plant

Algal O2-

tolerant 

Hydrogenase 

(1)

Cyanobacterium 

O2-tolerant 

Hydrogenase (2) 

Algal 

Sulfate 

Permease

Immobilized 

Algal, 

Sulfur 

deprived

PNS 

bacteria

B-1 B-2 B-3 B-4 B-5

Modules 1 1 1 1 1

Raceways/

Module 20 20 38 68 54 

Raceways/

Worker 100 100 100 100 100 

Shifts/Day 3 3 3 3 3

Workers 

FTE63 3 3 3 3 3 

Winder 

Workers N/A N/A N/A 0.5 FTE N/A

Total FTE 3 3 3 3.75 3

6.3.3 Energy Usage

The usage of utilities, feedstocks and creation of byproducts can be found on the 

Input_Sheet_Template worksheet of the model. There are no H2A default values for these. 

We are not aware of any Byproducts associated with the bio-hydrogen pathways. In the 

case of utilities, the utility (electricity, natural gas, steam) of interest is selected from a dropdown box. The amount consumed is based on the plant design and thus varies. For our 

1,000kgH2/day system those quantities are shown in Figure 6-6. The model has a cost rate 

for each utility and thus computes the total costs of utilities. 

Figure 6-6. Utilities Usage

Algal O2-

tolerant 

Hydrogenase 

(1)

Cyanobacterium 

O2-tolerant 

Hydrogenase (2) 

Algal 

Sulfate 

Permease

Immobilized 

Algal, Sulfur 

deprived

PNS 

bacteria

B-1 B-2 B-3 B-4 B-5

Electricity 

(kWh/kgH2) 4.36 4.36 6.22 2.26 4.81 

Water 

(gal/kgH2) 2.900 2.900 2.852 2.864 1.479 

In our bio-hydrogen pathways the feedstocks are nutrients and CO2 which are not H2A 

options that can be selected from the drop-down box. Therefore the usage, price and 

heating value of each feedstock must be entered manually. Those pathway specific required 

inputs are shown in Figure 6-7. 

 

63 Full Time Equivalent 
87 

Figure 6-7. Feedstock Usage

6.3.4 Specified Yearly Replacement Costs

The specified yearly replacement costs can be found on the Replacement Costs worksheet of 

the model. There is no H2A default value for this. This analysis recognizes that the 

transparent film which serves to capture product gas, permit sunlight, and keep culture free 

of contamination will need periodic replacement. The film is made of LDPE and over time 

will degrade such that less sunlight will enter the system thereby lowering plant efficiency. 

This analysis assumes that a 5-yr replacement cycle for the LDPE film is sufficient to keep 

system performing at acceptable levels. All other components are anticipated to operate for 

20 years. There are no other specified replacement costs over the twenty year analysis 

period for any of the pathways. Figure 6-8 lists the pathway specific costs associated with 

replacement components. 

Figure 6-8. LDPE Film Replacement Costs

Algal O2-

tolerant 

Hydrogenase 

(1)

Cyanobacterium 

O2-tolerant 

Hydrogenase (2) 

Algal 

Sulfate 

Permease

Immobilized 

Algal, Sulfur 

deprived

PNS 

bacteria

B-1 B-2 B-3 B-4 B-5

Quantity of 

Film64 87,080 (m2) 87,080 165,452 381,075 235,116 

Costs 

(yrs. 5, 10, 

15)

$60,932 $60,932 $115,770 $205,113 $164,516 

6.3.5 Baseline Uninstalled Costs

The baseline uninstalled costs can be found on the Capital Costs worksheet of the model. 

There is no H2A default value for this. These are the capital costs of the equipment that 

were computed in an earlier part of this project. Those separately calculated values will be 

imported into this area of the H2A Production Model to access the economic benefit of the 

capital expenditure. These costs are provided in the bill of materials for each pathway.

 

64 Quantity includes pathway area and overages explained in Section 5.1.2. 

Value Unit B1 B2 B3 B4

Lower Heating Value mmBTU/ Short Ton 9.37 9.37 9.37 9.37

Price Startup Year $/Dry Short Ton $ 1,400 $ 1,400 $ 1,400 $ 1,400

Startup CO2 lb 5,241 5,241 41,629 206,180

Usage lb/day 566 - 566 -

Usage DryTon/kg H2 2.83E-04 2.83E-04 0.00E+00 0.00E+00

Value Unit B1 B2 B3 B4 B5

Lower Heating Value mmBTU/ Short Ton 28.44 28.44 28.44 28.44 28.44

Price Startup Year $/Dry Short Ton 400 400 400 400 400

Usage DryTon/kg H2 1.16E-07 1.16E-07 2.16E-07 5.03E-07 3.09E-07

Carbon Dioxide

Nutrients
88 

6.3.6 Installation Cost Factor

The installation cost factor parameter can be found on the Capital Costs worksheet of the 

model. The default value for this is 1.3, however there is an exception. The installation of 

the reactor beds doesn’t rely on a cost factor but rather performs a separate excavation cost 

calculation which was described above. 

6.4 Results for Levelized Hydrogen Costs

The results from the H2A model are provided in tabular form. The total cost of produced 

hydrogen in a 10 tonne per day (TPD) plant65

Figure 6-9

is provided as well as the cost breakdown of 

that value over several cost components. Our analysis indicates that for a 10 TPD plant the 

B-1 and B-2 pathways produce the lowest cost hydrogen as shown in . B-1 and 

B-2 provide identical cost of hydrogen because the key assumptions effecting cost are the 

same for these two pathways. The primary components of the levelized cost of hydrogen 

are capital costs, fixed O&M and other variable costs as shown in Figure 6-10. The 

discussion section of this report will delve further into the costs included in these 

components. 

Figure 6-9. Levelized costs for B-1 through B-5 Pathways 

Cost Component

System B1 B2 B3 B4 B5

Capital Costs $1.74 $1.74 $2.82 $3.87 $3.37

Decommissioning Costs $0.02 $0.02 $0.03 $0.04 $0.03

Fixed O&M $0.60 $0.60 $0.99 $1.99 $1.27

Feedstock Costs $0.40 $0.40 $0.00 $0.00 $5.43

Other Raw Material Costs $0.00 $0.00 $0.00 $0.00 $0.00

Byproduct Credits $0.00 $0.00 $0.00 $0.00 $0.00

Other Variable Costs 

(including utilities) $0.23 $0.23 $0.33 $0.12 $0.25

Total $2.99 $2.99 $4.17 $6.02 $10.36

Total Cost of Produced H 2

Hydrogen Production Cost Contribution ($/kg)

 

65 Cost results are based on a 10 metric ton per day hydrogen plant composed of ten 1 metric ton per day 

modules.
89 

Figure 6-10. Comparison of Levelized Cost Components 

$0.00

$2.00

$4.00

$6.00

$8.00

$10.00

$12.00

B1 B2 B3 B4 B5

Other Variable Costs 

(including utilities)

Feedstock Costs

Fixed O&M

Decommissioning 

Costs

Capital Costs

6.5 Near Term Performance Results

In addition to calculating the results for the upper bound solar-to-hydrogen (STH) 

conversion efficiencies, which has been explicated thus far in the report, we also calculated 

levelized hydrogen costs for “near term” STH efficiencies. 

The upper bound performance projections for the B-1, B-2, B-3, and B-5 assume that there 

is full utilization of the solar photons for hydrogen production due to development of 

mutants that have reduced antennas and do not have saturation and other electron transfer 

limits. For near term organisms, utilization will be significantly less than 100%, but is 

expected to be better than the current status. Based on these anticipated developments, we 

have made estimates of the organisms’ near term properties and estimated the consequent 

near term production capabilities. 

Assumed near term solar-to-hydrogen (STH) conversion efficiencies are shown in Figure 

6-11 along with the previously discussed upper bound efficiencies for convenient 

comparison. . For B-4, the upper bound solar utilization comes from the NREL 

extrapolation of experimental results to generate a model of the best possible future 

performance level of B-4. For the B-4 near term projection, we have assumed a reduction 

in utilization from that best future performance upper bound. 

Levelized H2 costs corresponding to the near term efficiencies are shown in Figure 6-12. 

The primary components of the levelized cost of hydrogen are capital costs, fixed O&M, 

and “Other Variable Costs”. 

Figure 6-11. Near Term STH Efficiency Assumptions

STH Efficiency B1 B2 B3 B4 B5 

Upper Bound 9.2% 9.2% 5.2 % 2.25% 3.5%

Near Term 2% 2% 1.3% 1.5% 1.5%
90 

Figure 6-12. Levelized H2 Costs based on Near Term STH Efficiency

Cost Component

System B1 B2 B3 B4 B5

Capital Costs $5.61 $5.61 $7.04 $5.36 $7.10

Decommissioning Costs $0.06 $0.06 $0.07 $0.05 $0.07

Fixed O&M $2.08 $2.08 $2.96 $2.92 $2.62

Feedstock Costs $0.00 $0.00 $0.00 $0.00 $3.73

Other Raw Material Costs $0.00 $0.00 $0.00 $0.00 $0.00

Byproduct Credits $0.00 $0.00 $0.00 $0.00 $0.00

Other Variable Costs 

(including utilities) $0.41 $0.41 $0.41 $0.11 $0.44

Total $8.15 $8.15 $10.48 $8.44 $13.95

Total Cost of Produced H 2

Hydrogen Production Cost Contribution ($/kg)

Figure 6-13. Comparison of Levelized Cost Components for Near Term Efficiencies

$0.00

$2.00

$4.00

$6.00

$8.00

$10.00

$12.00

$14.00

$16.00

B1 B2 B3 B4 B5

Other Variable Costs 

(including utilities)

Feedstock Costs

Fixed O&M

Decommissioning 

Costs

Capital Costs

7.

The results discussion will look at how plant sizing impacts the key costs components of 

levelized costs, namely, capital costs, fixed O&M, and variable costs. An explanation for the 

low feedstock costs is also provided. 

Discussion of Results

This report’s baseline cost analysis assumes a 10 TPD plant consisting of ten 1 TPD 

modules. One would expect that the levelized costs results from H2A would be different if 

the plant size chosen was different. Plant sizes of 1TPD, 50TPD and 100TPD were 

reviewed and our analysis shown below indicates that the 10TPD facility was appropriate 

for this analysis. 

7.1 Capital Costs

Because this analysis is developing a large plant size from several smaller modules, one 

would expect equipment cost advantages based on modularity. However, recall that each 1 
91 

TPD module already has multiple raceways resulting in large quantities of similar 

components within a single module. Additional capital cost reductions as a result of 

increased purchase quantity are likely to be minimal since the quantity of most materials is 

already substantial for one module. The smallest module considered amongst our five 

pathways has twenty raceways, and the costing already employs a significant economy of 

scale. Further scaling of lower quantity, larger items, such as drum filters could result in 

further cost reduction, but has not been considered in this report primarily because low 

yearly sales of these items makes scaling data unavailable. 

Additionally, most components selected are available commercially off the shelf (COTS) 

and costs were obtained using vendor quotations. Because of this, the capital costs for each 

module already take into account commodity; high volume pricing and thus capital costs 

due to large plant size are expected to increase linearly. Figure 7-1 demonstrates the 

relationship between capital costs and plant size.

Figure 7-1. Capital Cost Increases from Increased Plant Size66

$-

$50,000,000 

$100,000,000 

$150,000,000 

$200,000,000 

$250,000,000 

$300,000,000 

$350,000,000 

$400,000,000 

$450,000,000 

0 Tonne 20 Tonne 40 Tonne 60 Tonne 80 Tonne 100 Tonne 120 Tonne

Capital Cost

Plant Size

Capital Costs of Increased Plant Size

B1

B2

B3

B4

B5

Exceptions to this COTS and high volume logic are the compressors and PSAs that are sized 

for 1TPD modules. Given the extremely large areas of a single module, we chose not to 

combine several modules into a common compressor and PSA due to the difficulty of 

transporting low pressure gas over long distances. Furthermore, unlike the other equipment 

whose costs were obtained from vendor quotes, the cost of compressors and PSAs are 

derived from H2A unit costing models. However those models already have built in 

assumptions to provide an nth unit cost so again no economy of scale can be assumed for 

these components. 

Because the capital costs do not achieve economy of scales we explored fixed O&M for an 

indication of optimal plant size.

 

66B1 and B2 are the exact same and overlap on the graph, which is why B1 does not show up on the graph 
92 

7.2 Fixed O&M (Plant Labor)

After capital cost component, fixed O&M is the next largest cost component of the levelized 

hydrogen cost. Within H2A the fixed O&M consists of primary of labor, G&A (which is a 

percentage of labor), Property Taxes (which is a percentage of capital costs), and Production 

Maintenance and Repair (which is a percentage of capital costs). Since we have already 

established that capital costs scale linearly, we focus on plant labor costs. 

One of the most significant differences in plant sizes is the amount of labor needed. Labor 

rates for a 1, 10, 50, and 100 TPD plant were examined. A summary of the assumed labor 

requirements and rates is shown in Figure 7-2. The analysis assumes that the plant will need 

constant supervision (3shifts/day) from workers who can fix any equipment malfunction and 

monitor plant operation. We assume one worker can effectively watch over 100 raceways 

working a standard 8-hour shift. Additionally, for the multiple module plants, an overall 

supervisor will be needed. For the 100 TPD plant, an assistant supervisor is added due to 

the extremely large area of such the plant. The B-4 system requires an extra worker in order 

to change one roll of immobilized algae film per day. 

Figure 7-2. Labor Assumptions for Different Plant Sizes

1 ton 10 ton 50 ton 100 ton 1 ton 10 ton 50 ton 100 ton

Modules 1 10 50 100 1 10 50 100

Total Raceways 20 200 1000 2000 38 380 1900 3800

Raceways/Worker 100 100 100 100 100 100 100 100

Shifts/Day 3 3 3 3 3 3 3 3

Workers FTE 3 6 30 60 3 12 57 114

Roll/Unroll Workers N/A N/A N/A N/A N/A N/A N/A N/A

Supervisor FTE 0 3 3 3 0 3 3 3

Asst. Super. FTE 0 0 0 1 0 0 0 1

Total FTE 3 9 33 64 3 15 60 118

FTE/1TPD Module 3 0.9 0.66 0.64 3 1.5 1.2 1.18

1 ton 10 ton 50 ton 100 ton 1 ton 10 ton 50 ton 100 ton

Modules 1 10 50 100 1 10 50 100

Total Raceways 90 900 4500 9000 54 540 2700 5400

Raceways/Worker 100 100 100 100 100 100 100 100

Shifts/Day 3 3 3 3 3 3 3 3

Workers FTE 3 27 135 270 3 18 81 162

Roll/Unroll Workers 0.8 8 38 75 N/A N/A N/A N/A

Supervisor FTE 0 3 3 3 0 3 3 3

Asst. Super. FTE 0 0 0 1 0 0 0 1

Total FTE 4 38 176 349 3 21 84 166

FTE/1TPD Module 3.8 3.8 3.5 3.5 3 2.1 1.68 1.66

B4 B5

B1/B2 B3
93 

When these plant sizes are analyzed in H2A, hydrogen costs decrease with plant size as 

shown in Figure 7-3. This study indicates labor cost reductions are achieved through 

increasing plant size. In all cases most of the cost benefit is gained in the initial increase in 

plant size, with the curve flattening towards a 100TPD system. Therefore, it is appropriate 

to choose the 10TPD plant for analysis. This is also intuitive from the labor assumptions in 

Figure 7-2 where a comparison in Total Raceways and Raceways/Worker shows that in the 

1 ton module workers are underutilized. 

Figure 7-3. Labor Cost Reductions from Increased Plant Size

$2.00 

$3.00 

$4.00 

$5.00 

$6.00 

$7.00 

$8.00 

$9.00 

$10.00 

$11.00 

0 Tonne 20 Tonne 40 Tonne 60 Tonne 80 Tonne 100 Tonne

$/kg H2

Plant Size

Labor Cost Reductions

B1

B2

B3

B4

B5

7.3 Variable Costs

The third largest cost component of the levelized hydrogen cost is variable cost. Within 

H2A the variable costs consists of utilities, byproduct credits, waste costs, and tax 

incentives. In our analysis only utilities (specifically electricity) are considered. Since 

electricity consumption is based on the equipment and we have chosen to duplicate modules 

and therefore equipment in the larger plants, variable costs are expected to be constant per 

kgH2 produced regardless of plant size. 

7.4 Feedstock Costs

Feedstock Costs are essentially the cost of nutrients. As we can see they are minimal in all 

pathways except B-5, where the cost of acetate is substantial. Since nutrient quantity is 

based on algae and area then we expect the relationship with plant size to be linear. Nutrient 

costs are currently derived from price quotes for high-quantity agriculture fertilizer 

purchases and data on international market prices for acetate. Considering the low level of 

nutrients consumed compared to conventional agriculture, change in pricing due to 

increased plant size is negligible and ignored. 
94 

7.5 Reactor Footprint

The above cost analysis uses a baseline 10 TPD plant consisting of ten 1 TPD modules. 

Thus all systems produce the same amount of hydrogen. However, due to varying 

efficiencies, the five systems result in vastly different land area requirements. If we were to 

normalize based on the land area, or reactor footprint, for the 1 TPD module we can analyze 

the production data in a different way. Choosing the area of the B-4 systems which has the 

largest reactor footprint, in Figure 7-4 we show the hydrogen outputs of each system if 

scaled to match the land area of B-4. This is an important consideration in areas where land 

may be in short supply. 

Figure 7-4. Normalized Hydrogen Production 

System B-1 B-2 B-3 B-4 B-5

Reactor Size (m2) 352,070 352,070 352,070 352,070 352,070

Algae Removed

(kg/day) 175 175 NA NA 468 

Hydrogen 

Production 

(kgH2/day)

4,348 4,348 2,331 1,000 1,628 

8.

Under the premise that low cost hydrogen is the most important factor, then the H2A 

analysis points us to B-1 and B-2 pathways as the best photobiological pathways. However, 

one engineering concern for B-1 and B-2, which produce stoichiometric mixtures of 

hydrogen and oxygen, is to assure that the gases are isolated from any potential ignition 

sources. Before moving forward with any given system further, additional sensitivities 

should be done on key assumptions in this analysis. In the previous section we explored the 

assumptions of plant size and modular design. One parameter that could be further tested is 

the impact of plant efficiency on levelized costs. The characterization section of this report 

identified 3 efficiencies for each system: maximum theoretical, assumed reactor operation, 

and current experimental data. This baseline analysis is based on the assumed reactor 

operation. However if that value cannot be reached, the impact on cost can be great. 

Another parameter to consider is component lifetime, which impact replacement costs. 

Sensitivities on labor rates for plant personnel could also be performed. Furthermore, the 

probability of success and commercialization timeline of all systems should be considered 

prior to eliminating any pathway from further analysis. 

Photobiological System Conclusions and Recommendations 

As mentioned in Section 6.4, a large portion of the levelized cost is from the capital costs of 

the plant. Figure 8-1 shows the contribution of each subassembly to the overall plant costs. 

The most costly subsystems are the Photobioreactor and Control System. 
95 

Figure 8-1. Capital Cost Allocations by Subassembly

42%

34%

4%

13%

7%

0%

B1

42%

34%

4%

13%

7%

0%

B2

49%

20%

18%

5% 7%

1% 0%

B3

67%

13%

2%

1% 12%

3% 2%

B4

61% 12%

2%

16%

9%

0% 0%

B5

Photobioreactor Subassembly

Gas Capture Subassembly

Organism Feed Subassembly

Recycle Subassembly

Control System

Initial Consumables

Roll-Unroll System

Within the photobioreactor subassembly the largest cost contributors are the pond 

installation and paddlewheel mixers. The pond installation is directly proportional to the 

number of raceways. The costs of the installation were well vetted as described in Section 

5.2.5. The plant location could affect labor costs associated with installation and could 

“provide some relief to capital cost. The second largest cost contributors to this 

subassembly are the paddlewheel mixers. These costs are well known. The quantities 

required are conservative estimates based on the land area and water depth. Further research 

could be done in this area to see if more cost effective mixing solutions are available for our 

area and bed depth. Any other solutions should have a similar lifetime to be cost effective 

or significantly cheaper to offset replacement costs.

The control system subassembly has two items which are of large costs. These are the gas 

analyzers and level indicators. Rather than placing a gas analyzer in each raceway H2 outlet 

pipe, we group the outlet of two raceways into one analyzer to reduce cost. If the sensor 

indicates a problem with the gas steam, a worker would need to identify which of the two 

raceways is affected using a handheld gas analyzer: this approach is considerably less 

expensive. An important consideration with this equipment is that each raceway contains 

one of these items thus a cost savings in a single unit is magnified by the number of 

raceways. Currently the lifetime on all this instrumentation is equal to the plant life so any 

substitute equipment should have similar lifetimes or take into account replacement costs. 

Rather than assume that each raceway is instrumented, a separate analysis could consider 

these instruments as portable. In that way there would be 1 of each sensor per worker rather 

than per raceway. This method would rely more heavily on the plant personnel which may 

diminish the number of raceways that can be monitored per worker. This tradeoff could be 

considered as a separate sensitivity study.
96 

9. Part 1 Appendix A: Key Terms and Definitions

Solar-to-Hydrogen (STH) Efficiency (in %): This is the ratio of the Lower Heating Value 

(LHV) of the net hydrogen produced from the system divided by the total solar energy input 

striking the biological beds. Hydrogen LHV is defined to be 33.33kWh/kg67. Solar input is 

defined as total photon energy striking the bed measured in W/m2

. We typically use daily 

averages and thus base efficiency calculations on the average Southwest US insolation 

(average over the year) of 5.5kW/m2/day. 

Maximum Theoretical Solar-to-Hydrogen (STH) Efficiency (in %): Converting all of 

this 5.5kW/m2/day insolation into electrons used for H2 production yields a 12.2% STH 

efficiency. Thus our nominal 9.2% STH is close to the theoretical maximum. The 12.2% 

max efficiency is based on:

1000W/m2 full sunlight intensity with 44% PAR (photosynthetically active radiation) and an 

average wavelength of 550 nm (average of 400-700nm) 

 => Photon density of 2.024x10-3 ein/m2

/s hitting the surface of the water.

Since it takes 4 photons to make an H2, the maximum energy of H2 produced from 

conversion of all of the photons is: 

(2.024 x 10-3 ein/m2/s) x (mol H2/4 mol photons) x (2.016g H2/mol H2) x (kg H2/1000g H2) 

x (33.33 kWh H2/kg H2) x (3600s/hr) 

= 0.122 kWH2/m2

Taking the ratio of hydrogen energy to the photon energy yields the maximum Solar-toHydrogen efficiency: 

(0.122kWH2/m2

) / (1.0kWphoton/m2) = 12.2% 

Solar Peak Intensity and Daily Variation (W/m2): Based on actual daily intensity data 

with amplitude varied to achieve the target total solar daily insolation (5.5kWh/m2

/day). 

Doubling Time (in hours): The time, for an algal colony to double in concentration. 

Typically based on constant 1000W/m2 sunlight irradiance.

Hourly Algae Renewal Fraction (in %): This term is the reactor bed volume fraction that 

is removed every hour. Removal of algae is necessary to maintain a constant colony of 

living algae both to remove dead/unproductive algae and/or to remove new growth. It is 

assumed that all algae (dead, young, and old) are evenly mixed together and preferential 

removal is not possible. Thus all performance numbers are based on average algae 

performance.

Growth Reactor: A bioreactor with the main function of promoting algae growth. 

Production Reactor: A bioreactor with the main function of promoting hydrogen 

production. 

 

67 The HHV is defined to be 39.4 kWh/kgH2.
97 

Four types of Systems are analyzed:

Single Bed: In this cyclic bioreactor, a single bed first functions as a growth reactor 

(typically 1-4 days) and then switches to a production reactor (for 1+ day). The end of the 

production phase is determined by the H2 production rate dropping below a threshold limit 

due to death or other impairment of the algae. At the end of the production phase, the bed 

either returns to the growth phase to rejuvenate the algae or becomes partially or completely 

removed from the reactor and replaced with fresh, healthy algae.

Dual Bed: In the cyclic bioreactor system, a first bed serves as a growth reactor to supply 

algae to the production reactor second bed. After the algae in the production reactor reaches 

the end of its useful life, the algae is removed from the reactor.

Chemostat: Derived from Chemical environment is static, a chemostat system is one in 

which fresh medium is continuously added to the reactor and “old” reactor medium is 

continuously removed, so that the reactor volume remains constant. In this embodiment, 

two reactors are used, a growth reactor and a production reactor, but their operation is 

constant rather than cyclic. The growth reactor continuously produces new algae at a rate 

that is fed into the production reactor. The production reactor has the same mass rate 

removed to maintain the mass balance. A centrifuge separates the water from the “dead” 

algae with the water being recycled back to the growth reactor and the algae going to a 

landfill or fermentation unit.

Chemostat II: The Chemostat II is a single bed chemostat system in which growth and 

production occur simultaneously in the same reactor. In this yet unproven, idealized system, 

a portion of the incident photon energy (ideally 22%) is used for cell maintenance (i.e. to 

keep the algae cells unstressed and healthy) and the remaining photon energy (78%) is used 

for H2 production for a net Solar-to-Hydrogen efficiency of 9.2%. 

Algae Capture Fraction (in %): The mass capture fraction of algae that is separated by 

the filter (centrifuge or drum filters). This is nominally 99%. Once separated, the algae 

enter a fermentation unit or are sent to a landfill. 

Algae Water Content (in %): This refers to the water fraction of the algae after it is 

separated from the renewal fraction by the filters. This is nominally 90%. 
98 

10.

 B1 (daily) B2 (daily) B3 (3 days) B4 (2 days) B5 (daily) 

Start Conditions - H2 Produdction

Algae 1,622 1,622 74,187 98,821 4,330

Inputs - H2 Production

Process Water 10,985 10,985 38,807 22,367 5,614

Nutrients 0 - 0 - 9,113

CO2 257 257 - - -

Total 12,863 12,863 112,994 121,188 19,058

Outputs - H2 Production

H2 1,111 1,111 8,235 4,706 1,111

O2 9,005 9,005 - - -

H20 - - - - 168

H20 Vapor (Evaporation) 951 951 2,010 1,340 650

CO2 - - 68,696 1,999 606

Dissolved CO2 - - 21,173 49,355 11,725

Waste Algae 175 175 - - 468

Algae 1,622 1,622 12,879 63,789 4,330

Total 12,863 12,863 112,994 121,188 19,058

Balance 0.0 0.0 0.0 0.0 0.0

B1 B2 B3 (4 days) B4 (2 day) B5

Start Conditions - Growth

CO2 (Dissolved in water) 21,173 1,999

Algae 12,879 63,789

Inputs - Growth

CO2 (from PSA) 68,695 49,354

Process Water 39,477 22,367

Nutrients 1 1

Total 142,225 137,509

Outputs Growth

Algae 74,187 98,821

H2O Vapor (Evaporation) 2,680 1,340

O2 65,359 37,347

Total 142,225 137,509

Balance 0.0 0.3

B1 (2 days) B2 (2 Days) B3 (2 days) B4 (2 days) B5 (2 days)

Inputs - Initial Colony Growth

Algae - - - - -

Process Water 973 973 7,730 38,286 -

Nutrients 0 0 0 1 7,793

CO2 2,377 2,377 18,879 93,505 0

Total 3,350 3,350 26,610 131,793 7,793

Outputs - Initial Colony Growth

Algae 1,622 1,622 12,879 63,789 4,330

H2O - - - - 1,559

O2 1,728 1,728 13,730 68,004 1,904

Total 3,350 3,350 26,610 131,793 7,793

Balance 0.0 0.0 0.0 0.0 0.0

Mass Balance Growth Phase

Mass Balance H2 Production Phase

Mass Balance Initial Growth Phase

Part 1 Appendix B: Mass Balance 

Days of B3 Growth Mode 4

Days of B4 Growth Mode 1

Overall Growth of B3 in Full Growth Mode 5.76

Overall Growth of B4 in Full Growth Mode 1.55

Starting Concentration of Algae in B3 0.85 g/l

Starting Concentration of Algae in B4 1.81 g/l

End Concentration of Algae in B3 4.91 g/l

End Concentration of Algae in B4 2.80 g/l
99 

Part II: Algae Fermentative H2

Production Systems 
100 

11.

As mentioned previously in this report, a constant slip stream will be removed from the 

raceways in the B-2, B-2, and B-5 systems to ensure that the organism concentration 

remains static. There is no way to ensure that the organisms removed in the slip stream are 

all exhausted organisms, thus the slip stream volume will be a mix of old exhausted and new 

replete organisms. The organisms removed from this slip stream will from here on be 

referred to as waste organisms. These organisms, while spent of their photo-hydrogen 

capabilities can still produce hydrogen through a fermentation process. In order to obtain 

maximal H2 production, Part II of this report addresses use of the waste organisms as 

feedstock to a fermentation process using hydrogen-generating bacteria to produce 

additional H2. The process yields and economics will build on the analysis already 

completed for the Part I photobiological systems. 

Introduction

In Part I of this report, five photobiological hydrogen production systems were defined: 

three of the five pathways generate a stream of waste organisms suitable for subsequent 

fermentation. For the B-3 system, an increase in organism mass in the growth stage is 

needed to provide enough mass for consumption in the oxidative respiration needed to 

consume the oxygen produced by water splitting to keep the system anaerobic. This mass 

consumption through respiration will return the system back to its original concentration by 

the end of the production stage. Thus, no additional mass need be removed. 

For the B-4 system, the algae were immobilized on a non-fermentable polypropylene 

substrate to reduce cost. For this embodiment, feeding the waste organisms to a fermentor 

would be too costly and cumbersome, and thus the waste organisms from this system are not 

used for fermentation. In contrast, NREL research for immobilized algae utilizes an alginate

film for immobilization. Use of the fermentable alginate film in B-4 would increase B-4 

cost but also increase the potential C-4 fermentation H2 production, and that is an option that 

could be considered in the future. 

Subsequent to this fermentation process, liquid fermentation products, such as acetate, can 

be recycled into a PNS photobiological reactor bed, or they can be transferred to a Microbial 

Electrolysis Cell (MEC) for additional hydrogen-production. Figure 11-1 below shows 

process alternatives considered.
101 

Figure 11-1. Fermentation Process within a Biological System

12.

Fermentation is the process of deriving energy from the oxidation of organic compounds, 

such as carbohydrates. Dark fermentation reactions, as the name suggests, do not require 

light energy, so they are capable of constantly producing hydrogen from organic compounds 

throughout the day and night. In theory, any carbohydrate source could be used as a 

feedstock for fermentation. Algae and other biomass listed in 

Theoretical and Practical Fermentation Reactions

Figure 12-1 are representative 

fermentation feedstocks. 

Chlamydomonas reinhardtii or

Cyanobacteria Photobiological Reactor

H2 & O2

(when O2 tolerant)

H2, CO2

waste

organisms, water

Purple non-sulfur (PNS) bacteria

Photobiological Reactor

Fermentation

organisms,

bacteria, water

H2,

CO2

acetate,

other organic acids,

water 

nutrients, water

water

Chlamydomonas reinhardtii or

Cyanobacteria Photobiological Reactor

H2 & O2

(when O2 tolerant)

H2

CO2

waste

organisms, water

MEC

Fermentation

organisms,

bacteria, water H2,

CO2

acetate,

other organic acids,

water

nutrients, water

OR

lignocellulosic

biomass

(optional)

lignocellulosic

biomass

(optional)
102 

Figure 12-1. Potential Fermentative Feedstock Materials

Algae and other biomass

Corn stover hydrolyzate

Sugar cane juices

Wood fiber hydrolyzates from pulp mills 

Biomass hydrolyzates from agricultural resides

Energy crops

The feedstock is composed of a starch/glucose component as well as proteins and lipids and 

other components. If it were possible to convert all of the waste organism starch entering 

the fermentation processes to hydrogen, the stoichiometric result, using glucose as a 

representative of starch feedstock, would be:

C6H12O6 + 6H2O => 12H2 + 6CO2. 

However, this theoretical limit does not naturally occur68. Instead, the observed glucose 

reactions are a conversion to H2, CO2, and acetate which is represented as:

C6H12O6 + 2H2O => 4H2 + 2CO2 + 2C2H4O2. 

This is known as the Thauer limit and results in 4.5% H2 output by mass relative to the 

glucose component68. An alternative, less complete, reaction that also takes place forms H2, 

CO2, and butyl acid: 

C6H12O6 => 2H2 + 2CO2 + C4H8O2

In actual practice, a combination of these reactions takes place, yielding between 2 and 4 

moles of hydrogen per mole of glucose. 

13.

The organisms under consideration consist primarily of starches, lipids, and proteins. 

Multiple tests have been run by researchers with NREL support on fermentation of 

Chlamydomonas reinhardtii. The consortium ferments the starch and lipid and protein 

components, at least partially. However, that research is not yet complete

Algae Fermentation Tests

69. 

The most directly applicable data for starch fermentation is from Project 3.3, 

Photobiological Systems for Hydrogen Photoproduction, Task 3 – Integrated Systems. Test 

results from this project for FY2007 and FY2008 were provided as references for the design 

of the fermentation system. The First Quarter FY08 results are summarized in Figure 13-1. 

These results have a comprehensive compilation of algae components and hydrogen outputs 

and were used as guidelines for the fermentation analysis along with the more recent data 

from the emails referenced. 

 

68 Benemann, John R. and Paola M. Pedroni. 4.3 Biological production of H2: mechanisms and processes.

69 Maria Ghirardi. Project 3.3:Photobiological systems for Hydrogen Photoproduction. 1st Quarter 2008 

Report. 
103 

Figure 13-1. NREL Fermentation Experimental Data 70

Algae

Characteristics

Input mass - Algae

Dwell

time

[hr]

H2 Output 

Total

mass

[mg] 

Starch

mass

[mg] 

Starch

[% of

algae] 

lipid + 

protein

% 

H2

[μmol] 

weight %

per algae

[g/g] 

Mole %

H2 to

glucose

January 2008

sulfur deprived, 21 hr 5.77 1.87 32.4% NA 120

hr

max

17.68 0.618% 1.70

sulfur deprived,142 hr 29.8 1.79 6.0% NA 22.97 0.155% 2.31

sulfur and phosphate

deprived, 142 hrs

86.2 1.76 2.0% NA 38.44 0.090% 3.93

Results69 show that the extent of algal sulfur deprivation has a direct effect on the algae’s 

total glucan content (glucose, starch, and glycogen), with the percentage mass of starch in 

the tested algae varying from 32% at 21 hours of sulfur deprivation, to 6% at 142 hours of 

sulfur deprivation, to 2% at 142 hours of sulfur and phosphate deprivation. The bacteria 

used in fermentation tests with these algae samples carried out fermentation mainly on the 

starch component, so the H2 outputs were dependent on the starch percentages. For these 3 

samples, test results yielded H2 masses that were respectively 0.618%, 0.155%, and 0.090% 

of the algae cdw. The results from the tests referenced are described in Figure 13-1. Note 

that in the last case, the molar yield of H2 to glucose is 3.93, very close to the Thauer limit 

of 4.0, indicating that some H2 was generated from lipid and protein components. 

Previous work on the sulfur deprivation of C. reinhadtii71 indicated that the starch content of 

cells grown under sulfur replete conditions and resuspended in a sulfur-deprived medium 

(t=0) is similar to that of cells deprived of sulfur for 142 hours. For the purposes of this 

analysis, thus, we used the data from the second row of Figure 89 to represent the starch 

amount of cultures in the three photobiological systems. The data shown in the third row 

reflects the starch amount of cultures that are immobilized in alginate films (not addressed 

in this analysis).

Current NREL research with different bacteria consortia have shown that H2 can also be 

generated from the lipid and protein content of the algae. Therefore, it is expected that 

future research will raise H2 output from this algae, and it was assumed that an H2 output 

between the 0.62% level and the 0.16% level would be representative of future capabilities, 

and a level of 0.4% H2/algae cdw was used in this analysis. 

Though First Quarter FY08 results were for a test time of 120 hours, other tests (Fourth 

Quarter FY07 and Fourth Quarter FY08) indicate that cumulative H2 production plateaus 

after 72 hours into a fermentative cycle. Thus, it was assumed that the fermentation process 

took 72 hours. 

 

70 Maria Ghirardi. Project 3.3: Photobiological Systems for Hydrogen Photoproduction. 1st Quarter Report. 22 

January 2008.

71 Kosourov, S.; Seibert, M.; Ghirardi, M. L. (2003). Effects of Extracellular pH on the Metabolic Pathways in 

Sulfur-Deprived H2-Producing Chlamydomonas Reinhardtii Cultures. Plant and Cell Physiology. Vol. 44(2), 

2003; pp. 146-155; NREL Report No. JA-590-34437.
104 

13.1 Fermentor Feedstock

As discussed above, feedstock of interest is waste organisms from photobiological 

processes. To facilitate the integration system task that comes after this fermentation 

analysis, the feedstock quantities and composition evaluated here are those identified in the 

Photobiological Characterization Report.72 Based on the previous photobiological systems 

examined, three possible fermentor feedstocks are shown in Figure 89. These are the waste 

algae or bacteria from photobiological systems B-1, B-2, and B-5. For the sake of 

consistency, the fermentor systems will be tagged C-1, C-2, and C-5 to correspond with the 

photobiological system from which it obtains its feedstock. Because the efficiencies and 

composition of each of photobiological systems vary, the slurry feedstock to the fermentor 

will differ. Note that there is no C-3 or C-4 fermentation pathway for reasons discussed 

earlier. 

Figure 13-2. Analysis Feedstocks Defined (from 10TPD Photobiological System)

C-1 C-2 C-5 

Photobiological Reactor Bed 

Type Chemostat II Chemostat II Chemostat II

Organism Chlamydomonas 

reinhardtii

Cyanobacteri

a

purple non-sulfur 

proteobacterium (PNS)

Slurry extracted from 10 TPD

photobio reactor bed %/day 10% 10% 10%

(0.2 g/L organism density) gal/hr 96,391 96,391 257,414

L/hr 364,840 364,840 974,312

Organism mass extracted

(organism dry mass) kg/hr 73 73 195

kg/day 1,751 1,751 4,677

Fermentor Input

% organism/water by 

weight73,74 % 20% 20% 20% 

Mass organism / liter g/L 200 200 200

Daily Quantity L/day 8,756 8,756 23,383

Fermentation time hrs 72 72 72

days 3 3 3 

13.1.1 C-1 Feedstock

The organism in the C-1 system is an oxygen-tolerant hydrogenase, Chlamydomonas 

reinhardtii cc124 as defined in the B-1 Photobiological Pathway. These algae have been 

cultivated in a Chemostat II reactor bed75

 

72James, Brian D. Task B: Photobiological H2 Production Subsystems Characterization Report. NREL 

Contract# AFH-8-88601-01. 24 December 2008. 

which is continuously extracting a slipstream of 

organism/water slurry to maintain a constant cell concentration. As mentioned previously, 

since there is no way to sort the organisms so that only exhausted ones are expunged from 

73 Aden, A. et al.. Lignocellulosic Biomass to Ethanol Process Design and Economics Utilizing Co-Current 

Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. NREL Contract# DE-AC36-99-

GO10337. June 2002, NREL/TP-510-32438.

74 Remainder is water.

75 While a generic chemostat system in defined as a reactor in which a constant volume of medium is 

maintained, the Characterization Report defines a Chemostat II as a chemostat system in which organism 

growth and hydrogen production simultaneously occur. 
105 

the reactor bed, we must assume a normally distributed cell age within the slurry. The 

quantities listed in the table above are computed based on the slipstream amount from a 

standard reactor bed size. 

13.1.2 C-2 Feedstock

The organism in the C-2 system is cyanobacteria, an oxygen-tolerant hydrogenase, based on 

the Synechocystis PCC6803 mutant from the B-2 Photobiological Pathway. These 

organisms have also been cultivated in a Chemostat II reactor bed which is continuously 

extracting a slipstream of organism/water slurry to maintain a constant cell concentration. 

Certain filamentous cyanobacteria acquire both longer filament lengths with age and 

develop gas vacuoles with age leading to compulsory flotation, while younger less buoyant 

cells remain dispersed. Separation of organisms by age may thus be more readily carried out 

for these bacteria, allowing removal of only the old organisms. This attribute could be 

exploited in a future design 

The quantities listed in Figure 13-2 above are computed based on the slipstream amount 

from a standard reactor bed size. However there is no data on the fermentation of these 

organisms, thus we assume the organism fermentation is similar to Chlamydomonas 

reinhardtii. 

13.1.3 C-5 Feedstock

The organism in the C-5 system is a purple non-sulfur proteobacterium from the B-5 

Photobiological Pathway. These bacteria have been cultivated in a Chemostat II reactor bed 

which is continuously extracting a slipstream of organism/water slurry to maintain a 

constant cell concentration. Since there is no way to sort these organisms, we assume a 

normally distributed cell age. The quantities listed in the table above are computed based 

on the slipstream amount from a standard reactor bed size. 

14.

In 

Basic Fermentation System Diagram

Figure 11-1 the fermentation system within a biological process is shown. This section of 

the report will delve into that portion of the system and describe in detail the components 

and the operation. Figure 14-1 shows the fermentation system is made of two 

subassemblies, fermentation and gas capture. Each of these subassemblies is described 

further below.
106 

Figure 14-1. Fermentation System Design

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Fermentation

Vessels

From Rotary

Drum Filters

Inlet

Valves

FermentorSubassembly

Transfer

Pump

Outlet

Valves

 

14.1 Fermentation Operation Scheme

In a prior section, it was mentioned that H2 cumulative production at lab scale tests plateaus 

in a 72 hour cycle. Thus we have taken 72 hours to be the duration of the fermentation 

process. Data indicates that the production rate over the 72 hours is not constant, but varies 

from minimal output for the first 24 hours, to max rate over the next 44 hours, and drops 

down to zero over the next 4 hours. 

In the fermentation process at larger scales, mixing will be a major concern to remove 

oxygen and to control variables such as pH and residual oxidants (oxygen and nitrate). For 

this purpose, the fermentation tanks each include a tank agitator. In addition, it will be 

needed to continuously monitor and control variables such as pH. 
107 

Because of the 3 day fermentation period, there are 3 fermentors running in parallel, with a 

fourth in recharging status or in reserve. At the end of the 72 hour process the vessel will 

contain solids residue which will need to be cleaned out prior to beginning the next batch. 

The fourth vessel will be cleaned while the other three are used. 

There are two options available to handle the variability. First, downstream equipment can 

be sized for peak rates and operate at reduced capacity during non-peak production. Second, 

the downstream equipment can be sized for average rates, and the plant design and operation

adjusted to accommodate the variable gas production. The decision of which route to take is 

an economic one. In our case the downstream components are a compressor, gas separation 

system, and a flow meter. The flow meter operation is not impacted by the variability in 

hydrogen production, however the compressor and gas separator are. We choose to operate 

the downstream equipment at steady-state conditions and accommodate the variable gas 

production. 

There are several ways the plant design could be modified to ensure the gas capture 

subassembly encounters a constant and continuous flow of output gas. The most direct 

would be to install a buffer storage tank between the two subassemblies. While simple in 

practice, storage of hydrogen at low pressures is voluminous. Thus in our design we have 

opted to build several fermentation vessels and stage the start of the fermentation process in 

each so that the cumulative hydrogen production from all vessels is leveled. 

14.2 Fermentor Subassembly

The fermentor subassembly conveys algae into the fermentors and processes it at the 

necessary conditions to produce hydrogen. In order to accomplish this, the subassembly is 

comprised of a pump, inlet valves, fermentation vessels, outlet valves, and piping. The 

pump transfers wet organisms and bacteria into the subassembly. In an integrated system 

one could assume that the organisms are collected from the rotary drum filters described in 

the photobiological systems report. The inlet and outlet valves are considered manual and 

necessary to isolate each vessel. The need to isolate each vessel will be evident when the 

operation scheme is described. Piping is needed to move material into the vessels. Because 

the quantities are low, pipe cross-sectional area is also low. Only short piping runs are 

needed due to the relatively low diameter of the fermentation tanks and the resulting close 

spacing of the tanks. 

The last piece of equipment in this system is the fermentation vessel itself. These vessels 

are designed for atmospheric pressure and temperature.

15.

Based on the system configurations described above, bills of material (BOM’s) were 

prepared for each of the fermentative systems. These BOM’s appear in

Capital Cost of the Fermentation System

 Figure 15-1 through

Figure 15-2. 
108 

Figure 15-1: BOM for Fermentor Systems C-1 and C-2 (sized for the organism wastestream of a 10TPD B-1 or B-2 Photobiological Hydrogen production system)

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Fermentation Subassembly

Tank 2312 gal 1 gal $ 2.00 9,247 $ 18,493 This is a scaled cost based on the capital costs from NREL Report TP-510-32438

Slurry Piping - 0.5" 500 ft 1 ft $ 0.52 500 $ 260 Sized using Continuity Equation. Assumed a velocity of 1 ft/sec. 

Slurry Piping - 1" 3734 ft 1 ft $ 1.00 3,734 $ 3,734 Sized using Continuity Equation. Assumed a velocity of 1 ft/sec. 

Pump 2312 gal/day 4 gal/min $ 198.50 1 $ 199 http://www.grainger.com/Grainger/items/1P795

Inlet Valves 1 in 1 each $ 67.23 4 $ 269 http://www.valvestore.com/products.asp?dept=1485

Outlet Valves 0.5 in 1 each $ 46.85 4 $ 187 http://www.valvestore.com/products.asp?dept=1485

Water 18,493 gal 1 gal $ 0.001665 18,493 $ 30.79 H2A Costing

Gas Capture Subassembly

Compressor 0.29 kgmol/h 0.29 kgmol/hr $ 9,233.00 1 2,679 $ H2A Costing

PSA $ 5,738 1 $ 5,738 H2A Costing

Gas Piping - 0.5" 100 ft 1 ft $ 0.52 100 $ 52 Sized using the Continuity Equation. Assumes 1 mol of H2, 0.5 Mol of CO@, and water vapor

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

System Initial Cost $ 37,142

Figure 15-2: BOM for 10 TPD Fermentor System C-5 (sized for the PNS waste-stream 

of a 10TPD B-5 Photobiological Hydrogen production system) 

Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Fermentation Subassembly

Tank 6173 gal 1 gal $ 2.00 24,693 $ 49,386 This is a scaled cost based on the capital costs from NREL Report TP-510-32438

Slurry Piping - 0.5" 500 ft 1 ft $ 0.52 500 $ 260 Sized using Continuity Equation. Assumed a velocity of 1 ft/sec. 

Slurry Piping - 1" 5077 ft 1 ft $ 1.00 5,077 $ 5,077 Sized using Continuity Equation. Assumed a velocity of 1 ft/sec. 

Pump 6173 gal/day 4 gal/min $ 198.50 1 $ 199 http://www.grainger.com/Grainger/items/1P795

Inlet Valves 1 in 1 each $ 67.23 4 $ 269 http://www.valvestore.com/products.asp?dept=1485

Outlet Valves 0.5 in 1 each $ 46.85 4 $ 187 http://www.valvestore.com/products.asp?dept=1485

Water 49,386 gal 1 gal $ 0.001665 49,386 $ 82.23 H2A Costing

Gas Capture Subassembly

Compressor 0.77 kgmol/h 0.77 kgmol/hr $ 9,233.00 1 7,154 $ H2A Costing

PSA $ 8,115 1 $ 8,115 H2A Costing

Gas Piping - 0.5" 100 ft 1 ft $ 0.52 100 $ 52 Sized using the Continuity Equation. Assumes 1 mol of H2, 0.5 Mol of CO@, and water vapor

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

System Initial Cost $ 76,281

As shall be noted when examining the cost results, tank capital cost and labor assumptions 

play an important role in the cost examination. For this reason, they are discussed in more 

detail below. 

15.1 Tank Cost

Capital cost of the stainless steel fermentation vessels is based on data from an NREL 

report76

16.

by Aden et al that gives us a $2.00/gallon cost for fermentor tanks in this size range 

operating at atmospheric pressure. As shown in the system bills of material, typical 

fermentor tank size for systems C-1, C-2, and C-5 is around 2,000-7,000 gallons. . 

Based on the above assumptions, the daily hydrogen output resulting from fermentation of 

algae from the photobiological pathways (B-1, B-2 and B-5) was calculated.

Fermentation Outputs for C-1 through C-5 

 

76 Aden, A. et al.. Lignocellulosic Biomass to Ethanol Process Design and Economics Utilizing Co-Current 

Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. NREL Contract# DE-AC36-99-

GO10337. June 2002, NREL/TP-510-32438.
109 

Figure 16-1. Fermentative Hydrogen Ouput (from 10TPD Photobiological System)

C-1 C-2 C-5 

Photobiological 

Reactor Bed Type Chemostat II Chemostat II Chemostat II

Organism Chlamydomonas 

reinhardtii Cyanobacteria purple non-sulfur 

proteobacterium (PNS)

Hydrogen Produced kg/day 7.0 7.0 18.7

weight % relative to 

organism =0.4%

For the C-1/C-2 process this amounts to only 7.0 kg H2/day for the system producing 10,000 

kg/day from the photobiological process. 

17.

Thus far, this report has discussed the capital cost and expenditures associated with building 

a fermentative hydrogen plant. While this is critical and valuable information, the build 

decision is ultimately based on the financial and economic benefits of these capital 

expenditures. The investment is quite large and cannot be properly evaluated without some 

knowledge of the expected return on that investment. In order to evaluate the return, DTI 

has performed a discounted cash flow (DCF) analysis using the H2A Forecourt Production 

Model, Version 2.0. 

Levelized Costs Assumptions & Calculations (H2A)

Discounted cash flow is a method of evaluating an investment by estimating the present 

value of future cash flows and taking into consideration the time value of money. The 

method uses future free cash flow projections and discounts them using the weighted 

average cost of capital (WACC) to arrive at a present value (PV). Two basic DCF methods 

are the net present value (NPV) method and the internal rate of return (IRR) method, both of 

which take into account the time-value of money, and are similar to the methods used in 

computing interest-income on bank deposits. This analysis uses the IRR method. Recall 

that IRR is the discount rate where the PV of total cash inflows equals the PV of the total 

cash outflows. 

The H2A model uses the DCF methodology described above. However future cash inflows 

from sale of the hydrogen are highly speculative. Thus, the DCF results in this analysis are 

presented in a slightly different manner. Rather than predict cash inflows from sales and 

compute the IRR, our analysis pre-selects an IRR value which would be agreeable to 

investors. It then assumes sales are equal to the amount of product that can be produced 

over the analysis period. With those two values in places, this analysis computes the price 

of hydrogen necessary for the IRR at the given sales volume. This price is expressed in 

$/kgH2. This is often referred to as the levelized cost of hydrogen. Levelized cost is the PV 

of the all costs associated with building and operating a generating plant over its economic 

life, converted to equal payments. 

17.1 Standard H2A parameters

In order to develop levelized costs, several parameters must be defined. Because this 

analysis focuses on a plant which is still in its conceptual stage, many of the values for these 

parameters must be assumed. The assumptions are documented in this section. These are 

meant to represent a baseline system and analysis. Later in the discussion portion of this 

report, the assumed values can be altered for sensitivity analyses.
110 

Standard H2A values and assumptions shown in Figure 17-1 apply to all biohydrogen 

pathways. This list does not encompass all parameters which must be defined in order to 

run the analysis; just those were there is an H2A Default value which has been accepted for 

this analysis. The remaining parameters are defined later in this section. Some of those 

parameters are common to all fermentative pathways and others are pathways specific. No 

dispensing parameters are listed here because they are not factored into the analysis.

Figure 17-1. H2A Default Values and Assumptions used for all Fermentative 

Pathways77

Parameter Assumptions

 

77 Appendix 3: Default Values and Assumptions. H2A Production Model, Version 2.0 User Guide, Rev. 

DRAFT, June 2008, p. 63. 

Operating Period 20 years 

Burdened Labor Rate for Staff $50/hour for central model

CO2 Capture Credit Not included in base cases (default value = 0) 

CO2 Production Taxes Not included in base cases (default value = 0) 

Construction Period and Cash Flow 1 year modeled after forecourt 

Co-produced and Cogenerated 

Electricity Price 

$30/MWh 

Decommissioning 10% of initial capital for forecourt model 

Depreciation Type and Schedule for 

Initial Depreciable Capital Cost 

MACRS: 20 years for central model 

Facility Life 20 years for forecourt model 

G&A Rate 20% of the staff labor costs 

Hydrogen Pressure at Central Gate 300 psig

Hydrogen Purity 98% minimum; CO < 10 ppm, sulfur < 10 ppm 

Income Taxes 35% federal; 6% state; 38.9% effective 

Inflation Rate 1.9%, but with resultant price of hydrogen in 

reference year constant dollars 

Installation Cost Factor 1.3 for all models

Land Cost $5,000/acre purchased for central model; 

Licensing, Permits and fees $1000 for forecourt model 

O2 Credit Not included in base cases 

Production Maintenance & Repairs 1⁄2% of direct capital cost for central model

Property Taxes and Business 

Insurance 

2%/year of the total initial capital cost 

Reference Financial Structure 100% equity with 10% IRR; includes levelized 

hydrogen price plot for 0%–25% IRR; model 

allows debt financing 

Sales Tax Not included on basis that facilities and related 

purchases are wholesale and through a general 

contractor entity 

Salvage Value 10% of initial capital for central model 

Working Capital Rate 15% of the annual change in total operating costs 
111 

17.2 Pathway Common Parameters

In addition to the financial parameters defined by H2A in the previous section there are 

other inputs which must be quantified in order to carry out the DCF analysis. All inputs can 

be found on the following worksheets in the H2A model; 

• Input_Sheet_Template 

• ReplacementCosts

• CapitalCosts

Many of these parameters are specific to the location, operation, and type of plant. In the 

case of our biohydrogen pathways there are some parameters that are the same for all 

pathways and some that vary by pathway. The parameters in Figure 17-2 are common to all 

pathways. 

Figure 17-2. Parameters Common to all Pathways 

Parameter Assumed Value Worksheet Name

17.2.1 Operating Capacity Factor

The operating capacity factor can be found on the Input_Sheet_Template worksheet of the 

model. This analysis assumes that the plant and dispensing station have a 90% operating 

capacity. This capacity factor takes into considering things such as planned maintenance 

outages, forced outages, etc. 

17.2.2 Reference Year Dollars

The reference year dollars parameter is the year dollars in which the cost of hydrogen is 

reported. The H2A standard is to report out hydrogen costs in 2005 dollars. The model 

expects capital costs to be entered in 2005 dollars. In this analysis Reference Year 2005 

was selected.

17.2.3 Site Preparation

The site preparation parameter can be found on the Input_Sheet_Template worksheet of the 

model. In central plants, H2A defaults this value to 1% of direct costs. This analysis uses 

the same default value. 

 

Operating Capacity Factor 90% Input_Sheet_Template

Reference Year Dollars 2005 Input_Sheet_Template

Site Preparation 1% of direct costs minus 

raceway excavation 

costs 

Input_Sheet_Template

Engineering & design 1% of direct costs Input_Sheet_Template

Process Contingency 20% of direct costs Input_Sheet_Template

Project Contingency $0 Input_Sheet_Template

Up-Front Permitting Costs 0.5% of direct costs Input_Sheet_Template

Licensing, Permits, and Fees $1000 Input_Sheet_Template

Production Maintenance & 

Repairs

0.5% of direct costs Input_Sheet_Template
112 

17.2.4 Engineering & design

The engineering & design parameter can be found on the Input_Sheet_Template worksheet 

of the model. In central plants, H2A defaults this value to 13% of direct costs. This 

analysis uses the same default value. 

17.2.5 Process Contingency

The process contingency parameter can be found on the Input_Sheet_Template worksheet 

of the model. In central plants, H2A defaults this value to 15% of direct costs. This 

analysis uses 20% of direct costs due to uncertainties in the system configuration. 

17.2.6 Project Contingency

The process contingency parameter can be found on the Input_Sheet_Template worksheet 

of the model. In our analysis we have chosen to include all contingency factors in the 

process contingency parameter, thus the project contingency is set to $0. 

17.2.7 Up-Front Permitting Costs

The up-front permitting cost parameter can be found on the Input_Sheet_Template 

worksheet of the model. The H2A default for this parameter is 9% of direct costs. This 

analysis uses the same default value. 

17.3 Pathway Specific Parameters

The last type of parameters we identify are those which are specific to each pathway 

analyzed. Figure 17-3 lists these parameters and rules of thumb applied in computing their 

values. These are pathway specific because they are associated with feedstock, process 

design and plant design. 

Figure 17-3. Pathway Specific Parameters

Parameter Rule Applied Worksheet Name

17.3.1 Land Required

The land required parameter can be found on the Input_Sheet_Template worksheet of the 

model. In developing the plant design, a land area for the fermentor system was computed 

and then increased by 30% for conservatism. The 30% factor is meant to encompass area 

requirements for pump skids, compressors, gas and other process flow components and a 

small control room. The total land requirement for each pathway is shown in Figure 17-4. 

Land Required Input_Sheet_Template

Production facility plant staff Input_Sheet_Template

Utility Usage Electricity and water 

costs use H2A pricing

Input_Sheet_Template

Feedstock Usage None Input_Sheet_Template

Specified Yearly Replacement 

Costs

None ReplacementCosts
113 

Figure 17-4. Land Required for each Pathway

Pathway Area (m2)

C-1 1,012

C-2 1,012

C-5 1,012

17.3.2 Production facility plant staff

The production facility plant staff parameter can be found on the Input_Sheet_Template 

worksheet of the model. This represents the number of full-time employees required to 

operate the plant. There is no H2A default value for this. Our analysis for each of the 

pathways is shown in Figure 17-5. 

Figure 17-5. Plant Staff Requirements for Fermentor plant

Pathway Total Workers Needed

C-1 1 man/3 shifts/day

C-2 1 man/3 shifts/day

C-5 1 man/3 shifts/day

17.3.3 Energy Usage

The usage of utilities, feedstocks and creation of byproducts can be found on the 

Input_Sheet_Template worksheet of the model. There are no H2A default values for these. 

Figure 17-6. Electricity Usage 

Pathway Electricity (kWh/kg H2)

C-1 6.15

C-2 6.15

C-5 3.90

17.3.4 Specified Yearly Replacement Costs

The specified yearly replacement costs can be found on the Replacement Costs worksheet of 

the model. There is no H2A default value for this and no fermentative system component 

replacements are assumed.

 

17.3.5 Baseline Uninstalled Costs

The baseline uninstalled costs can be found on the Capital Costs worksheet of the model. 

There is no H2A default value for this. These are the capital costs of the equipment that 

were computed in an earlier part of this project. Those separately calculated values will be 

imported into this area of the H2A Production Model to access the economic benefit of the 

capital expenditure. These costs are provided in the bill of materials for each pathway. 

17.3.6 Installation Cost Factor

The installation cost factor parameter can be found on the Capital Costs worksheet of the 

model. The default value for this is 1.3. 
114 

17.4 Levelized Costs

Once all the cost parameters are specified, the H2A Production Model can perform levelized 

cost calculations. Parameters can be changed in a sensitivity analysis to access the impact 

on the cost per kilogram of hydrogen. 

18.

Based on the system diagrams, performance, and capital cost presented above, the projected 

hydrogen cost was next calculated using the H2A analysis model

Cost Results for the Fermentative Systems

78

Figure 18-1

. The total levelized cost 

of hydrogen for each fermentative pathway is shown in . 

Figure 18-1: H2A Model Projected Hydrogen Cost per kg 

from Fermentor Systems C-1, C-2, and C-5 

Cost Component

System C1/C2 C5

Capital Costs $6.69 $2.91

Decommissioning Costs $0.03 $0.01

Fixed O&M $165.70 $62.10

Feedstock Costs $0.00 $0.00

Other Raw Material Costs $0.00 $0.00

Byproduct Credits $0.00 $0.00

Other Variable Costs (including utilities) $0.32 $1.14

Total $172.73 $66.17

Total Cost of Produced H 2

Hydrogen Production Cost Contribution ($/kg)

Note that the project cost of hydrogen is quite high (~$70-$170/kg) due to two primary 

factors. The first is the low kg/day of hydrogen production being produced from the system: 

7-19 kg H2/day. This has an amplifying affect since all costs must be amortized over a 

small amount of hydrogen. The second factor is labor cost, which is reflected under fixed 

O&M (operating and maintenance) cost. Labor cost is set at 1 laborer per shift, for each of 

the three shifts per day. Sensitivity to production scale and labor costs is examined in the 

next section. 

 

 

78 H2A Forecourt Production Model, Version 2.0.
115 

19.

For the 3 photobiological systems feeding the fermentor, each system producing 10 

tonnes/day, the major useful outputs from the fermentor are shown in 

Processing of Fermentor Liquid Outputs

Figure 19-1. 

Figure 19-1: Fermentor Outputs

Fermentor 

System

Fermentor 

Feedstock 

Source

Output from Fermentor (kg/day)

H2 CO2 Acetate Lipid,Protein

& Other

C-1, C-2 B-1, B-2 7.01 76.5 104.3 1563.6

C-5 B-5 18.7 204.1 278.6 4175.2

For the C-5 system, the acetate is used as feedstock to the ponds for PNS organism growth. 

It is not determined whether the acetate would be separated from the other fermentation 

byproducts or whether the entire byproduct stream could be used. At any rate, at the 

quantities produced in the fermentor, the acetate is only 3% of the B-5 system requirements. 

One alternative for obtaining additional gaseous products from the unused fermentation 

byproduct is a separate fermentation step. The remaining unreacted lipids and proteins 

could potentially be converted to gases (H2, CO2, CH4, and NH3) in additional fermentation 

reactions with different bacteria. Research work is ongoing to examine capabilities in those 

areas. 

Another post-processing alternative for the acetate and other simple organic liquids from C1 and C-2 is the use of a Microbial Electrolysis Cell (MEC) system. The MEC is an 

electrolysis unit that uses bacteria on the anode along with a moderate electrode voltage to 

carry out an electrolysis reaction of acetate (and other organic acids and alcohols) with 

water, producing H2 at the cathode and CO2 at the anode. The details of the MEC are 

discussed in a following section on lignocellulose fermentation and processing. 

The MEC sizing is determined by the volume of the diluted liquid organic input stream and 

the required residence time of the reaction. We assumed an acetate concentration of 2.0

gram acetate per liter. Residence time in the MEC depends on the voltage applied, with 

MEC size and cost decreasing with increased voltage up to a maximum of around 0.9-1.0 

volt. For these algae reactions a voltage of 0.9 Volts was assumed, which requires a 1 day 

residence time. For systems C-1 and C-2, (corresponding to biological pathways B-1 and B2), the processing of the liquid streams and the H2 output is shown in Figure 19-2. 
116 

Figure 19-2: MEC Flows

Parameters Units System C-1 and C-2

Acetate Flow Rate kg/day 104.3

Reactant water kg/day 56.3

MEC - H2 output kg/day 12.6

MEC Flow Rate gal/day 13,782

MEC Residence Time days 1

Total MEC Volume gal 13,782

Number of MECs tanks 1

Mass Balances summarizing the flows for C-1/C-2 and C-5 are shown in Figure 19-3 and 

Figure 19-4: 

Figure 19-3. C-1/C-2 Fermentation + MEC
117 

Figure 19-4. C-5 Fermentation

The heat balance will involve heating/cooling to maintain fermentation temperature. At the 

current stage of the algae fermentation research, this requirement has not been quantified.

20.

It is apparent that the initial fermentation step doesn't produce much H2 (~7-19 kg H2/day) 

when compared to the 10,000 kg H2/day photobiological system production rate from which 

it draws its organism feedstock. 

Algae Fermentation System Conclusions and Recommendations

If the full stoichiometric level of H2 fermentation of algal glucose were achievable through 

use of additional processing systems, one could potentially obtain three times more 

hydrogen. However, this is still only a small fraction of the targeted production rate of 

10,000 kg H2/day. It is also possible that a substantial amount of hydrogen can be obtained 

from the protein and lipid components, which constitute about 67% of the waste algae mass. 

Research is being carried out to optimize the fermentative process and to assess how much 

additional hydrogen might be available in these components. 

For the C-5 system, the fermentor byproduct acetate can be used as nutrient to the 

photobiological beds, providing 3% of the requirement. Future research might bring about a 

more complete fermentation reaction that would produce more acetate, more in line with the 

B-5 bed needs. 

The algal fermentor may also be useful to produce acetate for an MEC. Addition of the 

MEC provides an additional 180% H2, raising total C-1/C-2 output from 7kg/day to 20 

kg/day, but that is still a small amount compared with the photobiological production of 

10,000 kg/day. MEC operation is discussed in detail in the section on Lignolcellulosic 

fermentation and postprocessing. Integrated operation of a Fermentor and an MEC is 

discussed in the section on integrated system.
118 

Part III: Lignocellulosic Fermentative H2

Production Systems 

& Microbial Electrolysis Systems
119 

21.

An important pathway available for production of renewable hydrogen is fermentation of 

carbohydrate-rich feedstock. Fermentation is the process of oxidation of organic compounds, 

such as carbohydrates. Dark fermentation reactions with select bacteria are capable of 

producing hydrogen from organic compounds throughout the day and night. In theory, any 

carbohydrate source could be used as a feedstock. Examples of potential lignocelluloses 

fermentative feedstocks are shown in 

Lignocellulosic Hydrogen Production

Figure 21-1. The rate of H2 production via fermentation is 

relatively fast, and the reaction takes place continuously79. These features simplify reactor 

design and operation compared to photobiological pathways. 

Figure 21-1. Potential Fermentative Feedstock Materials

Algae and other biomass

Corn Stover

Switchgrass

Wood fiber 

The lignocellulosic biomass feedstocks of concern in this section of the report consist of 

cellulose, hemicellulose, and lignin components as shown in Figure 21-2. 

Figure 21-2. Lignocellulosic Biomass

Lignocellulosic

Biomass

Cellulose Hemicellulose Lignin

Long Polysaccharide 

chains consisting of a 

linear chain of D-glucose

Three monolignol 

monomers: (pcoumaryl, coniferyl, 

and sinapyl alcohols) are 

incorporated into lignin in 

the form of phydroxyphenyl, guaiacyl , 

and syringal

Shorter chains of a 

variety of sugar 

monomers including 

Glucose, Xylose, 

Mannose, Galactose, 

Rhammose, Arabinose, 

Etc....

 

79 Hawkes, F.R, Dinsdale, R., Hawkes, D.L., and Hussy, I. 2002. Sustainable fermentative hydrogen production: 

challenges for process utilization. Intl. J. Hydrogen Energy 27, 1339-1347.
120 

Several technical issues are concerns relative to the economic feasibility of the fermentation 

process. Those identified by the HFC&IT Program80 are:

• Glucose feedstock cost (barrier AT); 

• H2 molar yield (mol H2/mol hexose) (AR); and 

• Waste acid accumulation (AS)

 

The cost of feedstock and H2 molar yield thus play important roles in determining the viability 

of this process. This was confirmed by a preliminary boundary analysis for H2 production by 

fermentation, conducted by Tim Eggeman81 for a DOE Workshop on “Hydrogen Production via 

Direct Fermentation.” 

21.1 Fermentation Reactions

As cellulose and hemicellulose feedstock consist of long chains of glucose, xylose, and other 

sugar compounds, they can generally be represented in a reaction equation as glucose. If it were 

possible to convert all of the glucose in the fermentation processes to hydrogen, the 

stoichiometric result would be:

C6H12O6 + 6H2O => 12H2 + 6CO2. 

However, this theoretical limit does not occur in vivo because the overall energetics is not 

favorable. The complete set of enzymes needed for this multi-step conversion does not function 

in vivo toward hydrogen production. Instead microbes opt to produce a mixture of organic acids 

including acetic acid and butyric acid thus lowering the yield of hydrogen82. In an “ideal” 

fermentation reaction, the glucose reacts with water to produce H2 and CO2 gas plus acetic acid 

as the only non-gaseous by-product: 

C6H12O6 + 2H2O => 4H2 + 2CO2 + 2CH3COOH 

This is known as the Thauer limit and results in a maximal 4.5% H2 output by mass relative to 

the glucose component68. No known single microbe is able to achieve this yield. An alternative, 

less favorable, reaction could also take place and form H2 with butyrate as the organic acid 

product: 

C6H12O6 => 2H2 + 2CO2 + CH3CH2CH2COOH 

 

80 “Hydrogen, Fuel Cells & Infrastructure Technologies Program, Multi-Year Research, Development and 

Demonstration Plan,” Section 3.1 Hydrogen Production, U.S. Department of Energy: Energy Efficiency and 

Renewable Energy; Hydrogen, Fuel Cells, and Infrastructure Technology Program. April 2007. 

www.eere.energy.gov/hydrogenandfuelcells/mypp/

81 Eggeman, T. “Boundary Analysis for H2 Production by Fermentation, Neoterics International, Prepared for the 

National Renewable Energy Laboratory, March 12, 2004. Accessible from Proceedings of Workshop on 

“Hydrogen Production via Direct Fermentation.” 

http://www1/eere.energy.gov/hydrogenandfuelcells/wkshp_proceedings.html#hydrogen

82 Benemann, John R. and Paola M. Pedroni. 4.3 Biological production of H2: mechanisms and processes.
121 

To breakdown the cellulose and hemicelluloses, the fermentation process includes a hydrolysis 

step, a saccharification step, and, finally, a fermentation step for reacting the various sugars 

produced in the initial steps. In actual practice, the saccharification and fermentation reactions 

take place in the fermentor. In our analysis we assumed 90% efficiency of the hydrolysis step 

and 90% efficiency in the saccharification /fermentation step. This leads to 3.2 moles of 

hydrogen, 1.6 moles of carbon dioxide and byproducts consisting of organic acids and alcohols

plus unreacted glucose and other solids (lignin, xylane, etc.). This has been demonstrated in 

experiments. 

Production of 3.2 moles of hydrogen per mole of glucose leads to a mass conversion efficiency 

in the fermentation process of 3.6% (gram H2/gram cellulose and hemicellulose). The feedstock 

chosen for this analysis is corn stover which is only 64% cellulose and hemicellulose, leading to 

an estimated corn stover hydrogen conversion efficiency of 2.32% (gram H2/gram corn stover). 

21.2 Assumptions

While much of the process design for our hydrogen fermentation plant is derived from a 

Lignocellulosic ethanol report83, there are some important assumptions in the bacteria 

performance and process parameters that differentiate the two. 

21.2.1 Bacteria Assumptions

While testing for bacteria optimization is ongoing, we are hypothesizing performance based on 

hypothetical genetically engineered organisms capable of a fermentation H2 molar yield

approaching 4. We assume in our analysis that rates of H2 production are stable for long 

duration. There will be two types of organisms used concurrently that will carry out 

saccharification and fermentation of the cellulose and the hemicellulose, avoiding the need for 

separation and independent reactors. The organism that ferments the cellulose will create its 

own cellulase enzyme which obviates the need for an outside cellulase source. Finally, we are 

assuming that the non-gas products of the fermentation reaction are primarily acetic acid with 

lignins not reacting. 

21.2.2 Process Assumptions

Several of the processes involved in the ethanol fermentation report have been combined or 

eliminated in our version, simplifying the overall process. First, we assume that hydrolysis 

pretreatment breaks down the hemicellulose and cellulose inputs into reactable products. This 

pretreatment relies on high temperature acid hydrolysis alone, rather than the steam explosion 

process used in the ethanol report. We are using a feedstock cost that is minimal compared to a 

pure glucose feedstock. We have also assumed that saccharification and fermentation can occur 

in the same reactor. In our process, liming is done in the hydrolysis tank without any 

detrimental effect to the rest of the process. Lastly, we have assumed 20% solids content in the 

processed feedstock. 

 

83 Aden A. et al. Lignocelulosic Biomass to Ethanol Process Design and Economics Utilizing Co-Current Dilute 

Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. NREL/TP-510-32438, June 2002
122 

21.3 Lignocellulosic Fermentative Parameters

The baseline conditions that have been used for a fermentative pathway for hydrogen 

production using a lignocellulosic feedstock are summarized in Figure 21-3. 

Figure 21-3. Parameters for Fermentation System 

Note that in current tests the molar yield of H2 is closer to 2.0 mole H2 per mole of glucose. 

 

84 Assumes 15% water weight, for dry Corn Stover mass of 2,000 Metric Tonnes (MT)/day

85 Assumes 80% H2 recovery in the PSA. 

86 Volume is main fermentation tanks only- does not include pre-treatment tank volume.

System Parameters

Corn Stover Usage (MT/day) 235284

H2 Production Rate (kgH2/day)

 At Fermentor outlet

 At PSA outlet

46,477

37,18185

Plant Area (Acres) 11

Land Utilization (kg H2/acre/year) 1,112,422

Water Usage 

 (gal/day)

 (gal/kgH2) 

10,539

0.28

Electricity Required

 kWhr/day

 kWh/kgH2 

 kWh/kWh H2

152,389

4.10

0.12

Fermentor Parameters

Bacteria Cell lines Clostridium Consortium and Clostridium 

thermocellum

Duration of Cycle (hrs) 36

Target Temperature (oC) 55

Assumed Reaction Efficiency 

 Hydrolysis

 Saccharification and 

Fermentation

90%

90%

Feedstock - to - H2 Conversion 

Efficiency 

 gram H2/gram glucose 

 gram H2/gram corn stover

3.6%

2.3%

Reactor Volume86

 Gallons

 Liters

4,813,255

18,220,095

Theoretical Product Ratio 4 mol H2

2 mol CO2 

2 mol Acetate

Assumed Goal Product Ratio 3.2 mol H2

1.6 mol CO2 

1.6 mol Acetate
123 

21.4 Engineering Parameters

The lignocellulosic hydrogen production plant has been divided into 7 interconnected 

subassemblies, namely, 

• Corn Stover Prep Subassembly, 

• Pretreatment/Hydrolysis Subassembly,

• Fermentation Subassembly,

• Seed Production Subassembly, 

• Storage Subassembly, 

• Wastewater Treatment Subassembly, and

• Gas Compression and Separation Subassembly

A general diagram showing the basic flow of this system is shown in Figure 21-4. 

This diagram does not include all of the subassemblies. For a more detailed look at the 

interaction between the subassemblies please refer to the individual assembly diagrams shown 

in the subsequent sections. The subassemblies have several components that will also be 

identified in later sections of this document. 

Plant design was based upon a prior NREL report analyzing the design and costs of a 

lignocellulosic ethanol producing plant87. Using information from this report we extrapolated 

the costs and material requirements into a simplified, future hydrogen production plant using the 

same feedstock and similar processes. Hydrogen production involves different fermentative 

bacteria than ethanol production, and there are some additional assumptions made concerning 

potential advances in fermentation that will allow us to combine processes into one tank. These 

assumptions have been outlined previously in this report. Given the complexity of scaling such 

a complex system, we chose to emulate the ethanol plant’s scale. Therefore, all components in 

the following subassemblies are sized and priced analogous to the corresponding ethanol plant 

components. The volume of hydrogen produced is based upon the fermentation of 2,000 Metric 

Tons (MT) of corn stover (dry weight), which was determined by the ethanol report to be the 

most cost-effective volume of feedstock. 

Altering certain materials from the NREL report was considered for capital cost savings. 

Specifically, we analyzed use of concrete tanks instead of stainless steel. The Philadelphia 

Suburban Water Company conducted a life cycle analysis of concrete tanks and stainless steel 

tanks used in their water treatment system88. The conclusion drawn was that, the long-term 

durability of stainless steel made it the more cost effective tank solution. Coupled with the 

NREL reports decision to use stainless steel, we felt that there was no compelling reason to 

switch tank materials. 

 

87 Aden A. et al. Lignocelulosic Biomass to Ethanol Process Design and Economics Utilizing Co-Current Dilute 

Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. NREL/TP-510-32438, June 2002

88 Philadelphia Suburban Water Co. Steel vs. Concrete – A Life Cycle Analysis. 

http://www.steeltank.com/LinkClick.aspx?fileticket=L1jYUfta9aw%3D&tabid=219&mid=701
124 

Figure 21-4. Simplified Lignocellulosic Hydrogen Production Plant

Assumptions

1. Combine Saccharification and 

Fermentation Bacteria in single reactor

2. Use Bacteria that create their own 

Cellulase Enzyme

Hydrolysis 

Tank

80% H2 to Pipeline

System for Lignocellulosic Fermentation

Fermentor

1. Add 

Fermentation 

Bacteria

2. Fermentation of 

Hemicellulose

3. Cellulose 

Saccharifaction

4. Cellulose 

Fermentation

CO2, H2, 

H2O Vapor

Acetic Acid, Ethanol, and other Acids + Waste Lignin

Condenser

/HX Compressor PSA

H2O Liquid

CO2

Feed 

Shredder Wash Tank

Feedstock

Water

1. Acid hydrolysis of 

Hemicellulose

2. Ca(OH)2 to 

precipitate small 

soluble lignin

3. Neutralize Acid 

(Liming)

Heater 20% H2

H2O Vapor Water

Heat 

Exchanger

Air

Air

21.5 Corn Stover Prep Subassembly

The corn stover prep subassembly processes bales of corn stover in preparation for hydrolysis. 

The components of this subassembly in general terms are listed in Figure 21-5 and shown in

Figure 21-6. 

Figure 21-5. Corn Stover Prep Components

Organism Feed Subassembly Components

Bale Transport Conveyor 

Bale Unwrapping Conveyor

Belt Press Discharge Conveyor

Shredder Feed Conveyor

Truck Scales

Truck Unloading Forklift 

Bale Moving Forklift

Corn Stover Wash Table

Shredder

Concrete Feedstock-Storage Slab

Polymer Feed System

Wash Table Pump

Wash Water Pump
125 

Clarifier Underflow Pump

Clarified Water Pump

Belt Press Sump Pump

Clarifier Thickener

Belt Press

Magnetic Separator

Wash Water Tank

Clarifier Thickener Tank

Bales of corn stover are received by the plant on trucks. The trucks are weighed and unloaded 

by forklifts. Some bales are sent to storage while the rest are taken directly to conveyors. From 

there, the bales travel to an automatic unwrapping system that cuts away the plastic wrapping. 

Unwrapped bales are conveyed to a wash table, which starts breaking up bales and washes dirt 

from the corn stover. The washed stover is then conveyed past a magnetic separator which 

removes metal. Then, the stover is passed through primary and secondary shredders which 

shred the stover into a smaller, more easily processed size. Finally, the processed stover is 

conveyed to the pretreatment/hydrolysis subassembly. 

Dirty wash water is recycled and cleaned utilizing a clarifier-thickener system. The wash water 

is pumped to the clarifier where clean water is drawn off and recycled back to the wash tables. 

The underflow from the clarifier is then dewatered in a belt press. Because most of the wash 

water is recycled through this system, the fresh water requirement is low. 
126 

Figure 21-6. Corn Stover Prep Subassembly Design

Shredder Feed

Conveyor

Bale Unwrapping

Coneyor

Bale Transport

Conveyor

Impurities Waste Conveyor

Hay Bale Storage

Truck Scales

Shredder

To Hydrolysis

Corn Stover Wash Table

Wash Table

Pump

Wash Water

Pump

Polymer Feed

System

Clarifier

Thickener

Tank

Belt Press

Clarifier

Underflow Pump

Clarified Water

Pump

Belt Press

Sump Pump

Wash

Water

Tank

To Off-site

Magnetic Separator

The equipment in Figure 21-6 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $4,038,990 as shown in

Figure 21-7. 
127 

Figure 21-7. Capital Cost of Corn Stover Prep Subassembly 

Corn Stover Prep

Components

Material Chosen Total Pricing

Bale Transport 

Conveyor

400 ft long, 8' wide, 50 HP motor $ 800,000

Bale Unwrapping 

Conveyor

90 bale / hr

$ 300,000

Belt Press Discharge 

Conveyor

3' wide X 50' long.

$ 50,000

Shredder Feed 

Conveyor

8' wide X 30' long. 6001 lbs 31 fpm

$ 240,000

Truck Scales

Concrete. 8' wide X 30' long. 6001 lbs 31 

fpm $ 68,000

Truck Unloading 

Forklift

Propane Gas Forklift

$ 72,000

Bale Moving Forklift Propane Gas Forklift $ 72,000

Corn Stover Wash 

Table

A238. 55 ton/hr

$ 208,000

Shredder Alloy Steel. 55 ton/hr $ 1,208,000

Concrete FeedstockStorage SlabConcrete Slab. 350’ X538’

$ 450,655

Polymer Feed System $ 30,000

Wash Table Pump Rubber. 2500 gpm, 50 ft head $ 40,000

Wash Water Pump SS316. 5000 gpm, 50 ft head $ 30,000

Clarifier Underflow 

Pump

Rubber. MOD SRL 2X3-10 100 GPM --

50 TDH $ 6,000

Clarified Water Pump SS316. 5000 gpm, 50 ft head $ 15,000

Belt Press Sump Pump VJC 1.5X2.11 100 GPM -- 40 TDH $ 19,000

Clarifier Thickener 5000 GPM mechanism $ 135,000

Belt Press 304 SS. 1.5 meter $ 100,000

Magnetic Separator Tramp iron magnet separator $ 10,355

Wash Water Tank (20' Diameter X 22' Tall, 50,000 gal) $ 50,000

Clarifier Thickener 

Tank

Cement. (80' diameter) $ 135,000

TOTAL $4,038,990

21.6 Pretreatment/Hydrolysis Subassembly

The feedstock corn stover consists primarily of three components: cellulose, hemicelluloses, 

and lignin. Based on the Aden et al. report we have assumed a composition by weight of 36.7%

cellulose, 27.4% hemicellulose, 30.8% lignin, and 5.1% other non-reacting matter. The 

feedstock is mixed with water to the ratio of 20% dry weight and 80% liquid. The purpose of 

the pretreatment/hydrolysis subassembly is to break down the components of the corn stover 

into complex sugars on which bacteria are capable of carrying out saccharification and 
128 

fermentation. The components of this subassembly in general terms are listed in Figure 21-8

and shown in Figure 21-9. 

Figure 21-8. Pretreatment/Hydrolysis Components

Pretreatment/Hydrolysis Subassembly 

Components

Hydrozylate Mix Tank Agitator

Hydrozylate Mixing Tank

In-line Sulfuric Acid Mixer

Sulfuric Acid Pump

Sulfuric Acid Tank

Lime Solids Feeder

Lime Unloading Blower

Lime Dust Vent Baghouse

Lime Storage Bin

Hydrolysis Tank Pump

Hydrolysis Feed Boiler

The pretreatment/hydrolysis subassembly utilizes dilute acid hydrolysis reactions at 150 oC and 

4.7 atmospheres to convert the washed and cleaned corn stover feedstock into hemicellulosederived soluble sugars and crystalline cellulose that allow for subsequent bacterial 

sacchrification and fermentation. Cleaned corn stover from the corn stover prep subassembly 

mixed with water is heated by two heat exchangers, HX-1 and HX-2, and then heated in a boiler 

to 150 o C. In the boiler, heat is added to the mixture by burning the byproductH2/CO2 mixture 

coming from the waste side of the PSA separator. The heated slurry is then pumped into the 

hydrolysis reactor. The hydrolysis reaction hydrolyzes and breaks down the hemicelluloses 

portion of the feedstock by adding dilute (1.1%) sulfuric acid at 150 oC. The slurry is treated in 

this way for two minutes before the overliming process begins. This process raises the pH and 

re-solidifies the soluble lignin components for later separation. The subassembly components 

are shown in Figure 21-9. 
129 

Figure 21-9. Pretreatment/Hydrolysis Subassembly Design

Hydrolysis Tank

Agitator

Hydrolysis Tank

Pump

Hydrolysis Tank

Lime Storage

Bin

Lime Dust Vent

Baghouse

Lime Solids Feeder

Lime

Unloading

Blower

Sulfuric

Acid Tank

Sulfuric Acid

Pump

To Fermentation

Subassembly

Inline Sulfuric

Acid Mixer

From Corn

Stover Prep

Heat

Exchanger #1

(Fermentation

Subassembly)

Heat

Exchanger #2

(Fermentation

Subassembly)

Hydrolysis

Feed Boiler

The equipment in Figure 21-9 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $847,017 as shown in 

Figure 21-10. 
130 

Figure 21-10. Capital Costs of Pretreatment/Hydrolysis Subassembly

Pretreatment/Hydrolysis 

Components

Material Chosen Total Pricing

Hydrozylate Mix Tank

Agitator

SS. Top-Mounted, 1800 rpm, 50 hp

$ 28,421

Hydrozylate Mixing Tank SS304. 24770 gal., 13' dia. X 25' high, 15 

min. res. time, 90% wv, atmospheric $ 32,237

In-line Sulfuric Acid Mixer SS304. Static Mixer, 248 gpm total flow. $ 2,542

Sulfuric Acid Pump SS304. 4 gpm, 245 ft. head $ 8,289

Sulfuric Acid Tank Plastic. 6444 gal., 24 hr. residence time, 

90% wv $ 9,441

Lime Solids Feeder A285C. 8" dia., 140 cfh, 7000 lb/hr max 

flow conveyor $ 3,900

Lime Unloading Blower C.S. 7425 cfm, 6 psi, 22275 lb/hr $ 99,594

Lime Dust Vent Baghouse A285C, Polyester. 8333 cfm, 1389 sf, 6 

cfm/sf $ 140,707

Lime Storage Bin 4455 cf, 14' dia x 25' high, 1.5x rail car vol., 

atmospheric $ 136,370

Hydrolysis Tank Pump SS304. 737 gpm, 200 ft head $ 69,516

Hydrolysis Feed Boiler 6000 kW Boiler, 200 PSI. SS304 $316,000

TOTAL $847,017

21.7 Fermentation Subassembly

The fermentation subassembly carries out two processes: Saccharification and fermentation. 

The sacharification converts the hydrolyzer complex sugars into simple sugars using enzymes. 

Fermentation describes the subsequent conversion of these simple sugars into hydrogen gas by 

carefully designed fermentative organisms. As mentioned earlier, these organisms will produce 

a cellulase enzyme for the saccharification of the cellulose component of the feedstock. The 

components of this subassembly in general terms are listed in Figure 21-11 and shown in 

Figure 21-12. 

Figure 21-11. Fermentation Components

Fermentation Subassembly Components

Fermentation Tank

Fermentation Agitator

Fermentation Recirculation Pump

Lignin Wet Cake Screw

Acetate Product Storage Tank

Acetate Transfer Pump

Pneumapress Filter

Heat Exchanger #1

Heat Exchanger #2

Fermentor Temperature Control

In this subassembly the pretreated slurry is pumped into the saccharification/fermentation tank 

and inoculated with bacteria produced in the seed production subassembly, which will be 
131 

discussed later in this report. A fraction of the pretreated slurry will also be routed to the seed 

production subassembly for use in the cultivation of fermentative organisms. To maintain the 

temperature most conducive to fermentation, a portion of the slurry will be removed from the 

reactor, cycled through the fermentor temperature control heat exchanger for cooling, and 

returned to the tank. Residence time for one saccharification and fermentation cycle is a total of 

40 hours. 

In addition to hydrogen and carbon dioxide gases, the fermentation reaction will also produce 

acetic acid, which may be further processed into hydrogen by separate bacteria in a subsequent 

MEC system. This acetate is then pumped from the fermentation tank into temporary storage 

before being sent on for further hydrogen production or discarded. The subassembly 

components are all shown in Figure 21–12. 

Figure 21-12. Fermentation Subassembly Design

Fermentation

Tank Agitator

To Seed Production

Subassembly

To MEC

Fermentation

Recirculation Pump

Acetate Storage Tank

Temperature Control

Heat Exchanger

Innoculum from

Seed Production

Subassembly

Fermentation

Tank

Acetate Pump

From Hydrolysis

Gas Products to

Separation

Lignin Separator

To Hydrolysis - Heat

Exchanger

From Hydrolysis Heat

Exchanger

The equipment in Figure 21–12 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $8,458,507 as shown in

Figure 21–13. 
132 

Figure 21-13. Capital Costs of Fermentation Subassembly

Fermentation Components Material Chosen Total Pricing

Fermentation Tank SS304. 962,651 gal. each, 90% wv, API, 

atmospheric $ 2,466,955

Fermentation Agitator SS304. Side Mounted, 2 per vessel, 75 hp 

each, 0.15 hp/1000 gal $ 196,760

Fermentation Recirculation Pump SS304. 1060 gpm, 150 ft head $ 68,330

Lignin Wet Cake Screw 14" Dia X 100' Long $ 84,185

Acetate Product Storage Tank SS304. 456617 gal., 45' dia x 40' high, 4 hr 

res. Time, 90% wv, atmospheric $ 160,829

Acetate Transfer Pump SS304. 1231 gpm total, 615 gpm each, 100 

ft head $ 53,737

Fermentor Temperature Control Plate-Frame, SS304, atmospheric $ 33,905

Heat Exchanger #1 Plate-Frame, SS304, atmospheric $ 22,346

Heat Exchanger #2 Plate Frame, SS304, 4.68 atmospheres $ 194,905

TOTAL $ 8,424,607

21.8 Seed Production Subassembly

The purpose of the seed production subassembly is to breed hydrogen producing fermentative 

bacteria for use in the fermentation reactor. The components of this subassembly in general 

terms are listed in Figure 21-14 and shown in Figure 21-15. 

Figure 21-14. Seed Production Components

Seed Production Subassembly Components

Seed Hold Tank

Seed Hold Tank Agitator

Seed Hold Transfer Pump

1st Seed Fermentor

2nd Seed Fermentor

3rd Seed Fermentor

4th Seed Fermentor

5th Seed Fermentor

4th Seed Agitator

5th Seed Agitator

4th Seed Fermentor Coil

5th Seed Fermentor Coil

Seed Transfer Pump

The bacteria is grown in a series of seed fermentation vessels. Hydrolyzed slurry and nutrients, 

such as corn steep liquor (CSL) and Diammonium Phosphate (DAP) are combined with an 

initial seed inoculum grown in a very small vessel in a laboratory. The result of each seed batch 

is used as the inoculum for the next size seed increment. This series of scale-ups is continued 

until the last step is large enough to support fermentation. Finally, the seed inoculum, nutrients, 

and saccharified slurry are combined in each fermentation tank. The number of fermentors will 
133 

be five 1-million gallon vessels in a line with a 40 hour reaction residence time in each. The

subassembly components are shown in Figure 21-15. 

Figure 21-15. Seed Production Subassembly Design

3rd Seed

Fermentor

2nd Seed

Fermentor

1st Seed

Fermentor

Process

Water

DAP from

Storage

Subassembly

Corn Steep

Liquor (CSL)

from Storage

Subassembly

5th Seed

Fermentor and

Fermentor Coils

4th Seed

Fermentor and

FermentorCoils

Seed Hold

Tank

Seed

Transfer

Pump

Seed Hold

Transfer

Pump

Seed Hold

Tank Agitator

To Fermentor

5th Seed Fermentor

Tank Agitator 4th Seed Fermentor

Tank Agitator

Seed Production Subassembly

Slurry from

Hydrolysis

The equipment in Figure 21-15 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $1,007,829 as shown in 

Figure 21-16. 
134 

Figure 21-16. Capital Costs of Seed Production Subassembly

Seed Production 

Components

Material Chosen Total Pricing

Seed Hold Tank SS304. 233,333 gal., API atmospheric $ 160,829

Seed Hold Tank Agitator SS304. Top Mounted, 1800 rpm, 23 hp, 0.1 

hp/1000 gal $ 12,492

Seed Hold Transfer Pump SS304. 172 gpm, 150 ft head $ 22,050

1st Seed Fermentor SS304. 20 gal, jacketed, agitated, 1.3' dia., 

2' high, 15 psig $ 29,400

2nd Seed Fermentor SS304. 194 gal., jacketed, agitated, 3' dia., 

4' high, 2.5 psig $ 65,200

3rd Seed Fermentor SS304. 1950 gal., jacketed, agitated, 6.5' 

dia, 8' high, 2.5 psig $ 162,200

4th Seed Fermentor SS304. 19444 gal., 12' dia x 23' high, 

atmospheric $ 78,320

5th Seed Fermentor SS304. 194500 gal., API, atmospheric $ 293,097

4th Seed Agitator SS. Top Mounted, 1800 rpm, 6 hp, 0.3 

hp/1000 gal $ 23,289

5th Seed Agitator SS. Top Mounted, 1800 rpm, 19 hp, 0.1 

hp/1000 gal $ 20,582

4th Seed Fermentor Coil SS. 27 sf, 1" sch 40 pipe, 105 BTU/hr sf F $ 4,658

5th Seed Fermentor Coil SS. 307 sf, 3" sch 40 pipe, 92 BTU/hr sf F $ 28,238

Seed Transfer Pump SS304. 1231 gpm total, 615 gpm each, 100 

ft head $ 107,474

TOTAL $1,007,829

21.9 Storage Subassembly

The purpose of the storage subassembly is to provide bulk storage of chemicals used in the 

hydrogen production process. It also serves as the location for the main process water tank, 

which provides water needed in the various subassemblies of the system. The components of 

this subassembly in general terms are listed in Figure 21-17 and shown in Figure 21-18. 

 

Figure 21-17. Storage Components

Storage Subassembly Components

CSL Storage tank Agitator

CSL Pump

CSL Storage Tank

CSL/DAP Day Tank

CSL/DAP Pump

CSL/DAP Day Tank Agitator

DAP Solids Feeder

DAP Unloading Blower

DAP Vent Baghouse

DAP Storage Bin

Sulfuric Acid Pump

Sulfuric Acid Storage Tank
135 

Storage Subassembly Components

Propane Storage Tank

Cooling Water Pump

Make-up Water Pump

Process Water Circulating Pump

Process Water Tank

In addition to the main process water tank, stored chemicals include sulfuric acid, of which a 

five day supply is held, Corn Steep Liquor (CSL), a by-product of the corn wet-milling industry, 

and Diammonium Phosphate (DAP), both nutrient sources for fermentation seed growth. 

Propane gas for use in the bale handling forklifts. The subassembly components are all shown 

in Figure 21-18. 

Figure 21-18. Storage Subassembly Design

CSL/DAP Day

Tank Agitator

CSLTank

Agitator Sulfuric Acid

Storage Tank

Propane

Storage Tank

CSL Storage Tank

To Hydrolysis

To Forklifts

Sulfuric Acid

Pump

CSL/DAP

Pump

CSL Tank

Pump

CSL/DAP Day

Tank

To Seed Production

& Fermentation

Subassemblies

DAP Storage

Bin

DAP Vent

Baghouse

DAP Solids Feeder

DAP Unloading

Blower

Hopper

The equipment inFigure 21-18 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $924,749 as shown in

Figure 21-19. 
136 

Figure 21-19. Capital Costs of Storage Subassembly 

Storage Components Material Chosen Total 

Pricing

CSL Storage tank Agitator SS304. Top Mounted, 1800 rpm, 23 hp, 0.1 hp/1000 

gal

 

$ 2,143 

CSL Pump SS304. Top Mounted, 1800 rpm, 23 hp, 0.1 hp/1000 

gal

 

$ 6,188 

CSL Storage Tank Cs. 431 gpm, 150 ft head 

$ 61,949 

CSL/DAP Day Tank SS304. 10,000 gal, 24 hr res time, 90% wv, 9.5' dia 

x 18.9' high, atmospheric

 

$ 28,470 

CSL/DAP Pump CS. 431 gpm, 150 ft head 

$ 6,188

CSL/DAP Day Tank Agitator SS304. Top Mounted, 1800 rpm, 5 hp, 0.5 hp/1000 

gal

 

$ 12,348 

DAP Solids Feeder A285C. 8" dia., 140 cfh, 7000 lb/hr max flow 

$ 3,900 

DAP Unloading Blower CS. 7425 cfm, 6 psi, 22275 lb/hr 

$ 49,014 

DAP Vent Baghouse A285C, Polyester. 8333 cfm, 1389 sf, 6 cfm/sf 

$ 9,595 

DAP Storage Bin CS. 1425 cf, 9' dia x 19.5' high, 1.5x vessel vol. 

Req. for 7-day res time, atmospheric

 

$ 34,255 

Sulfuric Acid Pump SS316. 215 gpm, 150 ft head 

$ 13,814 

Sulfuric Acid Storage Tank SS316. 18697 gal, 120 hr res time, 90% wv, 12' dia 

x 22' high, atmospheric

 

$ 60,470 

Propane Storage Tank A515. 2000 gal, 10 day res time, 90% wv, 4' dia x 

23.6' length, 250 psig

 

$ 30,825 

Cooling Water Pump CS. 41000 gpm, 70 ft head. 

$ 242,216 

Make-up Water Pump CS. 1083 gpm, 75 ft. head 

$ 8,735 

Process Water Circulating Pump CS. 1199 gpm ea, 75 ft. head 

$ 17,635 

Process Water Tank CS. 756000 gal. 8 hr res time 

$ 337,004 

TOTAL $924,749

21.10Wastewater Treatment Subassembly

The function of the wastewater treatment subassembly is to treat process water for reuse to 

reduce the plant makeup water. The waste water is first screened, removing large particles to be 

subsequently sent to a landfill. Following screening, the organic matter in the water is then 

removed with anaerobic and aerobic digestion. This digestion produces a waste sludge and 

methane gas that could be used as a fuel for heating, however, we will not consider this further 

in this report. The components in this subassembly in general terms are listed in Figure 21-20. 
137 

Figure 21-20. Wastewater Treatment Subassembly Components

Wastewater Treatment Subassembly Components

Equalization Basin Agitator

Anaerobic Agitator

Anaerobic Digester Feed Cooler

Nutrient Feed System

Biogas Emergency Flare

Anaerobic Reactor Feed Pump

Aerobic Digester Feed Pump

Bar Screen

Equalization Basin

Anaerobic Digester

Aerobic Lagoon Agitator

Filter Precoat System

Aerobic Sludge recycle Pump

Aerobic Sludge Pump

Aerobic Digestion Outlet Pump

Sludge Filtrate Recycle Pump

Treated Water Pump

Belt Filter Press 

Aerobic Digester

Clarifier

The design for the wastewater treatment section mirrors that of the ethanol report, which relied 

upon the analysis of Merrick Engineering. Their design was carefully scaled and optimized for 

their flow rates, which differ slightly from ours. We have scaled linearly the components of the 

system to match our increased flow rate. The subassembly components are show in 

Figure 21-21. 
138 

Figure 21-21. Wastewater Treatment Subassembly Design

Bar Screen

Separator

Equalization Basin

Pump

Nutrient

Feed

Tank

Feed

Cooler

Anerobic

Digester

Pump

Aerobic

Digester

(Pond)

Clarifier

Filter

Precoat

Systems

Belt Filter

Press

Sludge Screw

Waste

Waste

Process Water

The equipment in Figure 21-21 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $12,039,997 as shown in

Figure 21-22. 
139 

Figure 21-22. Capital Cost of Wastewater Treatment Subassembly

Storage Components Material Chosen Total 

Pricing

Equalization Basin Agitator SS. 38 HP each, Fixed Prop, 0.1 HP/1000 gal 

$ 94,624

Anaerobic Agitator SS. Fixed Prop, 41 HP, 0.05 hp/1000 gal 

$ 100,955

Anaerobic Digester Feed 

Cooler SS316. TEMA BES type, floating head 

$ 428,475

Nutrient Feed System CS. 5 tanks and pumps 

$ 71,301

Biogas Emergency Flare SS. Flare and Pilot 

$ 69,099

Anaerobic Reactor Feed 

Pump CS. Centrifugal Pump, 876 GPM, 150 ft. head 

$ 37,983

Aerobic Digester Feed Pump CS. Centrifugal Pump, 830 GPM, 150 ft. head 

$ 35,651

Bar Screen CS. 0.5” Mehc. Cleaned Screen 

$ 392,551

Equalization Basin Concrete. 377516 Gal. Residence time 7.2 hr 

$ 1,168,809

Anaerobic Digester Expoxy Lined. 810250 gal each, space velocity 

12g/COD/day

 

$ 2,935,620

Aerobic Lagoon Agitator CS. Twister Surface Aerators, 50 HP each 

$ 1,665,920

Filter Precoat System Cs. Tank, Agitator, Pump 

$ 9,996

Aerobic Sludge recycle Pump SS316. Slurry Pump 2.5 GPM, 150 ft. head $ 36,983 

Aerobic Sludge Pump SS316. Slurry Pump 25.3 GPM, 150 ft. head $ 36,983 

Aerobic Digestion Outlet 

Pump CS. Centrifugal Pump, 828 GPM, 150 ft. head 

$ 35,651

Sludge Filtrate Recycle Pump CS. Centrifugal Pump, 22 GPM, 150 ft. head 

$ 20,324

Treated Water Pump CS. Centrifugal Pump, 803 GPM, 100 ft. head 

$ 35,317

Belt Filter Press 304SS. Filter Press Separator 

$ 2,166,438

Aerobic Digester Polymer Lined Pit. 19500000 gal, 16.3 day 

residence time. 

 

$ 2,116,294

Clarifier Concrete. 195,289 gal, residence time 3.9 hr 

$ 581,023

TOTAL $12,039,997

21.11Gas Compression and Separation Subassembly

The function of the gas compression and separation subassembly is to compress the gaseous 

outputs from the fermentor to 300 psi, separate the product hydrogen from water vapor and 

CO2, and deliver the hydrogen to the production facility limits. The outlet pressure of hydrogen 
140 

at the plant gate is 300psi. This pressure was selected to provide with a system comparable to 

other DOE H2A Production Plants. The components of this subassembly in general terms are 

listed in Figure 21-23. 

Figure 21-23. Gas Compression and Separation Components

Gas Capture Subassembly Components

Fermentor Compressor

Fermentor PSA Separator 

Fermentor Condenser 1

Fermentor Intercooler 1

Fermentor Intercooler 2

Hydrogen Flow Meter

The hydrogen/CO2/water vapor mixture is at atmospheric pressure and 55o

C at the outlet from 

the fermentor: 

• H2 46,477 kg/day 

• CO2 507,340 kg/day 

• H2O 113,542 kg/day 

The gas mixture is initially sent to a condenser to cool the gas to 45oC condense about 40% of 

the water. This counter-flow gas/water heat exchanger has been scaled from a comparable 

condenser used in the Ethanol Report. 

After the condenser, the gas enters a 2-stage compressor, which compresses the gas to 302 psi. 

to accommodate pressure losses in the PSA gas separation unit and the pipeline runs, which are 

estimated as 2 psi. A gas/water intercooler (FIC-1) is used between first and second stage 

compression to reduce the compressed gas temperature to 45oC and to condense additional 

water. A second intercooler (FIC-2) is used between second stage compression and the PSA to 

reduce the gas temperature to 45o

C and to condense additional water. In the condenser and 

intercoolers, a total of 97% of the water vapor from the fermentor output is condensed out 

before the fermentor PSA, with the remaining water vapor mole fraction reduced to 0.00460. 

The gas/water intercoolers were scaled from a comparable heat exchanger used in the Ethanol 

Report. The subassembly components are shown in Figure 21-24. 
141 

Figure 21-24. Gas Compression and Separation Subassembly Design

Hydrogen

Pipeline

Hydrogen

Flow Meter PSA

2 Stage

Compressor

Condenser

Intercoolers

The equipment in Figure 21-24 has been priced previously in several reports. Thus pricing for 

the equipment was easily located in reference documents and through discussions with 

suppliers. The capital cost of the subassembly components are shown in Figure 21-25. 

Figure 21-25. Capital Cost of Gas Compression and Separation Subassembly

Gas Compression and 

Separation Components

Material Chosen Pricing

Compressor Source: Using H2A Cost guidelines and scaling factors $15,729,564

PSA Separator Source: Using H2A Cost guidelines and scaling factors $1,527,494

Fermentor Condensor 1 Shell –Tube, SS304, atmospheric $ 29,623

Fermentor Intercooler 1 Shell-Tube, SS304, 4 atmospheres $ 57,651

Fermentor Intercooler 2 Shell-Tube, SS304, 20 atmospheres $ 311,057

Hydrogen Flow Meter Information from Emerson Process Management $5,000

Total $17,660,889

21.12Bill of Materials

The flow chart outlining mass flows and subsystem parameters is shown in Figure 21-26. The 

detailed mass and energy balance data is included in Part III, Appendix A. The full bill of 

materials for the lignocellulosic hydrogen production plant is shown in Figure 21-27. The 

system described in this section of the report has a capital cost of $44,944,078. 
142 

Figure 21-26. Lignocellulosic Fermentation Flow Chart 

Fermentation System - Batch Process

Efficiencies- Hydrolysis: 90%, Fermentation: 90% 

HX-2 10,000,000 kg/day HX-1 Feed Corn

 Heater 10,000,000 kg/day +85C pump +23C 10,000,000 kg/day Stock Stover

11,214 kW 128C 4.7 atm 43C 1 atm 20C Processors Feed Stock

 150C -85C -25C 20C

 4.7 atm 4.7 atm expander 1 atm 30C 20% solid Water

10,000,000 kg/day 9,332,641 kg/day Lignin 80% water Storage

Hydrolyzer Separator

Saccharification Liquid mash 1,732,783 kg/day

10,000,000 kg/day 10,000,000 kg/day HX 9,332,641 kg/day

 55C, 1 atm Fermentor

150C in Temperature acetate

140C out Control glucose 109,611 kg/day

4.7 atm Lime water 6,832,581 kg/day

Fermentation Gas:

10,000,000 kg/day H2 46,477 kg/day water

55C, 1 atm CO2 507,340 kg/day

H20 113,542 kg/day Byproduct Organic Liquids

Recovery

Condenser FC-1 Other

1 atm

14.7 psi liquid water

667,359 kg/day 49,066 kg/day

Fermentor Compressor Intercooler

stage 1 FIC-1

60 psi

618,293 kg/day 618,293 kg/day liquid water

Fermentor Compressor 45C 49,307 kg/day

stage 2 60 psi

302 psi

568,986 kg/day Intercooler H2 PSA High purity H2

 FIC-2 CO2 80% 37,181 kg/day

568,986 kg/day 556,692 kg/day CO2 507,340 kg/day

45C 40C H2O 2,876 kg/day Waste H2

302 psi H2 9,295 kg/day 301 kg/day

liquid water

12,294 kg/day

Waste H2

2,783 kg/day

 H2O

8,994 kg/day

with:

 CO2

657,666 kg/day

45C

490,873 kg/day

VH2O

Recovery

Water

Separation

and 
143 

Figure 21-27. Bill of Materials Lignocellulosic Fermentation Description Install Factor Size Req'd Units Unit cost Qty Req'd Total Cost

Corn Stover Prep Subassembly

Bale Transport Conveyor 1.62 $ 400,000 2 $ 800,000

Bale Unwrapping Conveyor 1.19 $ 150,000 2 $ 300,000

Belt Press Discharge Conveyor 1.89 $ 50,000 1 $ 50,000

Shredder Feed Conveyor 1.38 $ 60,000 4 $ 240,000

Truck Scales 2.47 $ 34,000 2 $ 68,000

Truck Unloading Forklift 1 $ 18,000 4 $ 72,000

Bale Moving Forklift 1 $ 18,000 4 $ 72,000

Corn Stover Wash Table 2.39 $ 104,000 2 $ 208,000

Shredder 1.38 $ 302,000 4 $ 1,208,000

Concrete Feedstock-Storage Slab 2.2 $ 450,655 1 $ 450,655

Polymer Feed System 2.28 $ 30,000 1 $ 30,000

Wash Table Pump 3.87 $ 20,000 2 $ 40,000

Wash Water Pump 5.19 $ 15,000 2 $ 30,000

Clarifier Underflow Pump 13.41 $ 6,000 1 $ 6,000

Clarified Water Pump 7.07 $ 15,000 1 $ 15,000

Belt Press Sump Pump 2.92 $ 19,000 1 $ 19,000

Clarifier Thickener 1.51 $ 135,000 1 $ 135,000

Belt Press 1.25 $ 100,000 1 $ 100,000

Magnetic Separator 1.3 $ 10,335 1 $ 10,335

Wash Water Tank 2.8 $ 50,000 1 $ 50,000

Clarifier Thickener Tank 3.04 $ 135,000 1 $ 135,000

Pretreatment/Hydrolysis Subassembly

Hydrozylate Mix Tank Agitator 1.2 $ 28,421 1 $ 28,421

Hydrozylate Mixing Tank 1.2 $ 32,237 1 $ 32,237

In-line Sulfuric Acid Mixer 1 $ 2,542 1 $ 2,542

Sulfuric Acid Pump 2.8 $ 8,289 1 $ 8,289

Sulfuric Acid Tank 1.4 $ 9,441 1 $ 9,441

Lime Solids Feeder 1.3 $ 3,900 1 $ 3,900

Lime Unloading Blower 1.4 $ 99,594 1 $ 99,594

Lime Dust Vent Baghouse 1.5 $ 140,707 1 $ 140,707

Lime Storage Bin 1.3 $ 136,370 1 $ 136,370

Hydrolysis Tank Pump 2.8 $ 69,516 1 $ 69,516

Boiler 1.3 $ 316,000 1 $ 316,000

Fermentation Subassembly

Fermentation Tank 1.2 $ 493,391 5 $ 2,466,955

Fermentation Agitator 1.2 $ 19,676 10 $ 196,760

Fermentation Recirculation Pump 2.8 $ 13,666 5 $ 68,330

Lignin Wet Cake Screw 1.4 $ 16,837 5 $ 84,185

Acetate Product Storage Tank 1.4 $ 160,829 1 $ 160,829

Acetate Transfer Pump 1.3 $ 53,737 1 $ 53,737

Pneumapress filter 1.4 $ 1,285,736 4 $ 5,142,942

HX - 1 1.3 $ 22,346 1 $ 22,346

HX - 2 1.3 $ 194,618 1 $ 194,618

FTC 1.3 $ 6,781 5 $ 33,905

Seed Production Subassembly

Seed Hold Tank 1.2 $ 160,829 1 $ 160,829

Seed Hold Tank Agitator 1.2 $ 12,492 1 $ 12,492

Seed Hold Transfer Pump 1.4 $ 22,050 1 $ 22,050

1st Seed Fermentor 2.8 $ 14,700 2 $ 29,400

2nd Seed Fermentor 2.8 $ 32,600 2 $ 65,200

3rd Seed Fermentor 2.8 $ 81,100 2 $ 162,200

4th Seed Fermentor 1.2 $ 39,160 2 $ 78,320

5th Seed Fermentor 1.2 $ 146,549 2 $ 293,097

4th Seed Agitator 1.2 $ 11,645 2 $ 23,289

5th Seed Agitator 1.2 $ 10,291 2 $ 20,582

4th Seed Fermentor Coil 1.2 $ 4,658 1 $ 4,658

5th Seed Fermentor Coil 1.2 $ 28,238 1 $ 28,238

Seed Transfer Pump 1.4 $ 53,737 2 $ 107,474

Storage Subassembly

CSL Storage tank Agitator 1.2 2,143 1 $ 2,143

CSL Pump 2.8 6,188 1 $ 6,188

CSL Storage Tank 1.4 61,949 1 $ 61,949

CSL/DAP Day Tank 1.4 28,470 1 $ 28,470

CSL/DAP Pump 2.8 6,188 1 $ 6,188

CSL/DAP Day Tank Agitator 1.2 12,348 1 $ 12,348

DAP Solids Feeder 1.3 3,900 1 $ 3,900

DAP Unloading Blower 1.4 49,014 1 $ 49,014

DAP Vent Baghouse 1.5 9,595 1 $ 9,595

DAP Storage Bin 1.3 34,255 1 $ 34,255

Sulfuric Acid Pump 2.8 13,814 1 $ 13,814

Sulfuric Acid Storage Tank 1.2 60,470 1 $ 60,470

Propane Storage Tank 1.4 30,825 1 $ 30,825

Cooling Water Pump/MEC Pumps 2.8 242,216 1 $ 242,216

Make-up Water Pump 2.8 8,735 1 $ 8,735

Process Water Circulating Pump 2.8 8,818 2 $ 17,635

Process Water Tank 1.4 168,502 2 $ 337,004

Wastewater Treatment Subassembly

Equalization Basin Agitator 1.2 28,400 3 $ 94,624

Anaerobic Agitator 1.2 30,300 3 $ 100,955

Anaerobic Digestor Feed Cooler 2.1 128,600 3 $ 428,475

Nutrient Feed System 2.58 21,400 3 $ 71,301

Biogas Emergency Flare 1.68 20,739 3 $ 69,099

Anaerobic Reactor Feed Pump 2.8 11,400 3 $ 37,983

Aerobic Digestor Feed Pump 2.8 10,700 3 $ 35,651

Bar Screen 1.2 117,818 3 $ 392,551

Equalization Basin 1.42 350,800 3 $ 1,168,809

Anaerobic Digestor 1.04 881,081 3 $ 2,935,620

Aerobic Lagoon Agitator 1.4 31,250 53 $ 1,665,920

Filter Precoat System 1.4 3,000 3 $ 9,996

Aerobic Sludge Recycle Pump 1.4 11,100 3 $ 36,983

Aerobic Sludge Pump 1.4 11,100 3 $ 36,983

Aerobic Digestion Outlet Pump 2.8 10,700 3 $ 35,651

Sludge Filtrate Recycle Pump 2.8 6,100 3 $ 20,324

Treated Water Pump 2.8 10,600 3 $ 35,317

Belt Filter Press 1.8 650,223 3 $ 2,166,438

Aerobic Digestor 1 635,173 3 $ 2,116,294

Clarifier 1.96 174,385 3 $ 581,023

Gas Capture Subassembly

Compressor Fermentor 1.3 1704 kgmol/hr $ 9,233 1,704 $ 15,729,564

FC-1 1.3 $ 29,623 1 $ 29,623

FIC -1 1.3 $ 57,651 1 $ 57,651

FIC -2 1.3 $ 311,057 1 $ 311,057

PSA Fermentor 1.3 1 $ 1,527,494 Hydrogen Flow Meter 1.3 $ 5,500 1 $ 5,500.00

System Initial Cost $ 44,944,078
144 

22.

A Microbial Electrolysis Cell (MEC), also referred to as a Bio-Electrochemically Assisted 

Microbial Reactor (BEAMR) can be used to generate H2 from simple organic liquids, such as 

acetic acid and ethanol, using electrolysis enhanced by a microbial reaction at the anode. The 

H2 gas is evolved at the cathode and CO2 gas is evolved at the anode. 

Microbial Electrolyis Cell (MEC) Hydrogen Production

In this section we will consider the MEC as a stand-alone hydrogen production unit, with a 

purchased acetic acid feedstock. Integration of the MEC system with the fermentation system 

to utilize the fermentator byproducts will be discussed in a later section. 

22.1 Theoretical Reaction and Assumptions

Assuming an acetic acid input and MEC electrode voltages above 0.4 Volts, a complete reaction 

in the MEC would be: 

C2H4O2 + 2H2O => 4H2 + 2CO2

The MEC process has been extensively researched and demonstrated at Penn State University. 

Output from a MEC using an acetic acid feedstock has been demonstrated in lab-scale tests to 

be as high as 96% + 1%89. With 95% reacted into H2 and CO2, 3.8 moles H2 is generated per 

mole acetic acid. For this reaction, the input acetic acid was diluted with water to 1 gram per 

liter. For our calculations we assumed an average of 90% efficiency due to scale-up losses and 

anode/cathode degredation. This leads to an estimated acetic acid to hydrogen conversion ratio

of 12.08% (gram H2/gram acetic acid).

The following major MEC assumptions are used in the analysis:

• Both the anode and cathode materials are low cost and sufficiently durable for long-term 

operation (20 years) 

• External electrical power will be provided to sustain the reactions. 

• The volume of acetic acid feedstock chosen for use was chosen based on the amount of 

acetic acid and glucose that would be used in an integrated system utilizing the 

fermentation effluent 

• The cost of the acetic acid feedstock is $0.595/kg based on the international market price 

as reported by the Korean Chemical Journal90

22.2 MEC Operating Parameters

For this study the baseline MEC system was sized so that it could be integrated with the 

fermentation system described in the previous section. The baseline conditions that have been 

used for MEC hydrogen production using an acetic acid feedstock are summarized in 

Figure 22-1. 

 

89 Call, Douglas and Logan, Bruce, “Hydrogen Production in a single Chamber Microbial Electrolysis Cell 

Lacking a Membrane.” Environmental Scientific Technology 42 (2008): 3401-3406. 

90 Lee, Sang Yup. “Plastic Bacteria? Progress and Prospects for Polyhydroxyalkanoate Production in Bacteria.” 

Trends in Biotechnology 14 (1996): 431-438. 
145 

Figure 22-1. Parameters for MEC System

 

91 Assumes 95% H2 recovery in the PSA.

92 Logan, Bruce et al. “Microbial Enectorlysis Cells for High Yield Hydrogen Gas Production from Organic 

Matter.” Environmental Scientific Technology 42 (2008): 8630-8640. 

93 Including electrode volume and spare capacity of 5%

MEC (0.9V)

Acetic Acid Usage (kg/day) 767,277

H2 Production Rate (kgH2/day) 88,08591

Plant Area (Acres) 10

Land Utilization (kg H2/acre/

year)

2,788,553

Water Usage 

 (gal/day) 

 (gal/kgH2)

9,242,555 

105

Electricity Required

 kWhr/day 

 kWh/kgH2 

 kWh/kWh H2

3,288,962 

37.34 

1.12

MEC Parameters

Bacteria Cell lines (anode) Pseudomonas spp. And Shewanella 

spp.92

Duration of Cycle (hrs) 24

Target Temperature (oC) 30

Assumed Reaction Efficiency 90%

Feedstock - to - H2 Conversion 

Ratio

 gram H2/gram acetic acid 12.08%

Reactant Volume

 Gallons 

 Liters

91,524,080 

346,682,121

Reactor Volume93

 Gallons 

 Liters

96,100,285 

364,016,227

Theoretical Product Ratio 4 mol H2

2 mol CO2

Assumed Actual Product Ratio 3.6 mol H2

1.8 mol CO2
146 

22.3 MEC System Description

The MEC hydrogen production plant has been divided into 3 interconnected subassemblies, 

namely, 

• MEC Subassembly,

• Storage Subassembly, and 

• Gas Compression and Separation Subassembly 

22.3.1 MEC Plant Layout

The interfaces between the subassemblies are shown in Figure 22-2. The subassemblies have 

several components that will be identified in later sections of this document. 

Figure 22-2. Simplified MEC Hydrogen Production Unit

Isolation Valves

MEC

Tank

Recirculation

Hydrogen Gas Line

1

Feed

Cascade Fill To MEC 2

Water/Acetate

Mixture

For the hydrogen production rate shown in the previous section, a large number of units, each 

comprised of a million gallon reactant tank and electrodes, are necessary. Residence time in the 

reactor at 0.9V potential is 24 hours per electrolysis cycle. Each reactant tank can be operated 

in a semi-continuous process, with daily replenishment of acetate and water. With the MEC 

system designed to consume 690,549 kg acetate and 414,314 kg water per day, these amounts 

will be replenished daily. Periodically, after an appropriate number of cycles, a tank would be 

emptied, cleaned and replenished with new electrolyte, and also anode organisms if necessary. 

A representative tank field layout is shown inFigure 22-3. 
147 

Figure 22-3. MEC Unit Field Layout

MEC Bank

~ 10 MEC and 1 Transfer Pump

MEC Module = ~ 12 MEC Banks

BANK

1

BANK

2

BANK

3

BANK

4

BANK

5

BANK

6

BANK

7

BANK

8

BANK

9

BANK

10

BANK

11

BANK

12

1 2 3 4 5 6 7 8 9 10

Water From Storage

Hydrogen to Gas Capture

Transfer

Pump

Waste

Water

Acetate From Storage

Mixing Valve

22.3.2 MEC Reactor Subassembly

As mentioned, to generate H2 from acetic acid or from the acetic acid, succinic acid, lactic acid, 

formic acid, glucose, and alcohol liquid byproducts of a fermentation reaction, a Microbial 

Electrolysis Cell (MEC) also referred to as a Bio-Electrochemically Assisted Microbial Reactor

(BEAMR), can be used to electrolyze the mix of organic materials and water to produce 

hydrogen. 

Ordinarily, a water electrolysis unit requires about 1.8 V to electrolyze water. The MEC uses 

microbe reactions at the anode with the organic materials, and water to generate a significant 

voltage. This is supplemented with an externally applied anode-to-cathode voltage to initiate 

and sustain an electrolysis process to generate H2 and CO2 from the organics and water. The 

current generated provides the electrons to produce H2 at the cathode, with CO2 produced at the 

anode. The microbes on the anode develop sufficient voltage such that the additional potential

needed for the electrolysis of organic byproduct is only 0.2-0.9 V. 

With acetic acid input, a complete redox reaction in the MEC would be: 

C2H4O2 + 2H2O => 4H2 + 2CO2 
148 

The components of the subassembly in general terms are listed in Figure 22-4. 

Figure 22-4. MEC Components

MEC Subassembly Components

MEC Tank

Cathode

Anode

Power Supply

Process Water Circulating Pump

MEC Transfer Pumps

MEC Intercooler 1

MEC Intercooler 2

22.3.2.1 MEC Cell Description

A representative MEC reactor cell is shown in Figure 22-5. Though the cell is shown with a 

PEM membrane, it can be operated without one, resulting in a small amount of gas dilution. 

Figure 22-5. MEC Cell

Tests have been run by PSU with a high surface area brush anode (brush fiber surface 

area/brush envelope volume = 18,200 m2

/m3) and a cathode made from carbon cloth loaded at

0.5 mg/cm2Pt94. Data from these tests in Figure 22-5 shows the linear relationship between H2

output Q (m3-hydrogen per day/m3 MEC) as a factor of applied voltage with Q=0 at V=0.2. 

The figure also shows energy efficiency increasing with decreasing voltage. However, at lower 

voltage, the reaction is slower and system size increases substantially. These tests run without a 

membrane provided good separation of the H2 and CO2. 

 

94 Call Douglas and Bruce E. Logan. “Hydrogen Production in a single Chamber Microbial Electrolysis Cell 

Lacking a Membrane.” Environmental Scientific Technology 42 (2008): 3401-3406.
149 

Figure 22-6: MEC Operating Performance – Pt Cathode95

Because of the high cost of Pt cathodes, additional research has been carried out by PSU, 

substituting a stainless steel brush for the carbon/Pt cathode96

Figure 22-6

. With a very high surface area 

brush cathode, the H2 production rate was only slightly less than with the Pt loaded cathode. 

These tests demonstrated a Q of 1.7 at a voltage of 0.6V. Assuming that the Q vs. V 

relationship for the brush cathode is also linear with a zero output at 0.2V, a line for the cell 

output through Q of 1.7 and V of 0.6 can be superimposed on . The resulting 

performance line is the solid blue line labeled “brush cathode” in Figure 22-7. The equation for 

the Pt Cathode Q is:

Q = 5.15V – 1.03 

And the similar equation for the Brush Cathode (the solid blue line in Figure 22-7) is: 

Q = 4.25V – 0.85 

 

95 Logan, Bruce et al. “Microbial Electrolysis Cells for High Yield Hydrogen Gas Production from Organic 

Matter.” Environmental Scientific Technology 42 (2008): 8630-8640. 

96 Call, Douglas et al. “High Surface Area Stainless Steel Brushes as Cathodes in Microbial Electrolysis Cells.” 

Environmental Scientific Technology (2008)
150 

Figure 22-7. MEC Operating Performance - Brush Cathode

The conclusion drawn from these performance graphs is that while a low voltage is more 

efficient from a power consumption standpoint, a high voltage results in a markedly smaller 

MEC volume for a given volume of H2 product, thus reducing capital costs significantly.

22.3.2.2 MEC Hydrogen Production Cost Optimization

MEC capital costs decrease with increased voltage, due to higher production rate. However, 

electrical costs increase with voltage. Thus, higher electrical costs go along with lower tank and 

electrode costs. These costs were modeled as a function of voltage to examine the minimum H2 

cost design for the MEC. The five primary H2 production cost components are: 

• MEC Tank

• Brush Cathode 

• Brush Anode97

• Power Supply, and 

, 

• Electric Power

Figure 22-8 plots the cumulative H2 production cost buildup from these five components. We 

used component sizing methodology used by PSU. The most dominant factor was the cost of 

stainless steel brush wires (based on SS wire costs from various suppliers). Though we looked 

at wire diameters down at 0.02 mm, the best cost-to-surface-area ratio was with 0.08 mm wire. 

To reduce brush costs, we reduced the surface area required by a factor of 2 relative to the 

 

97 Logan, Bruce et al. “Graphite Fiber Brush Anodes for Increased Power Production in Air Cathode Microbial 

Fuel Cells.” Environmental Scientific Technology (2007) 
151 

brushes used in the PSU tests, per the suggestion of Bruce Logan who pointed out that the brush 

size is not yet optimized. 

The results shown in Figure 22-8 indicate that 0.9V achieves the lowest cost. We therefore used 

this point in our cost analyses. 

Figure 22-8. MEC Cost Optimization

0

1

2

3

4

5

6

7

8

9

10

0.4 0.5 0.6 0.7 0.8 0.9 1.0

$/kg Hydrogen

Electrode Potential (Volts)

MEC Hydrogen Production Cost

plus power supply

plus electric power

plus tank

plus anode

cathode cost

22.3.2.3 MEC Brush Cathode

For the brush cathode, the ratio of its surface area to the brush envelope volume, S/Ve, was 

approximately 2500 m2/m3. For the PSU test cells, the ratio of brush cathode surface area to the 

reactor volume, S/Vr, was between 650 and 810 m2/m3. A top view of the MEC cell with a 

half-brush anode and half-brush cathode which gave the best performance is shown in Figure 

22-9. The brushes are semicircular to improve the packing density and reduce the ion transport 

distance relative to circular brushes. 
152 

Figure 22-9: MEC Test Unit Top view with Half-Brush Cathode and Anode

For the MEC design and costing analysis, carbon fiber brush anodes and 304 stainless steel 

brush cathodes as discussed in the references were assumed. Cell performance was based on 

the test results but reduced from 95% to 90% to account for scale-up losses and brush 

degradation. The specific brush sizes used in the PSU tests were laboratory scale (1 inch length 

by 0.5 inch radius) and are obviously not appropriate for large-scale production cells. For this 

analysis, the surface area ratio used was reduced to half of that used in the small scale tests. 

Using this ratio for a large system the materials requirements are quite substantial, and effort 

will have to be dedicated in the future to designing lower cost high effective area cathodes and 

anodes amenable to a large scale system. 

In addition, to reduce costs, the cells were operated without an ion exchange membrane. 

Potential future developments future for a scaled-up MEC system would include: 

• reduced reactor volume per hydrogen generation, increased reactant concentration 

• lower cost cathodes

• additional bacteria-generated voltage 

• minimized CH4 output 

• optimized solution electrolyte & conductivity

• addressal of long term corrosion issues 

• cathode and anode scale-up for large cells 

22.3.2.4 Total Hydrogen Production Cost

The system size was chosen to be compatible with the fermentation system output reactable 

organic liquids. For the stand-alone MEC, a similar mass flow of acetate was substituted for the 

output organic liquid mass flow of the fermentor. Thus, the feedstock reacted for the standalone MEC was:

• Acetic Acid: 767,277 kg/day (12,788 moles/day) 
153 

For the overall system layout, the output H2 gas from the multiple MEC units feed a manifold 

that feeds a single gas compression and separation subassembly, discussed in the following 

section.

For this effluent, the capital cost of the subassembly is $541,705,673 as shown in Figure 22-10. 

Figure 22-10. Capital Costs of MEC Subassembly

Seed Production 

Components

Material Chosen Total Pricing

MEC Tanks SS304. 962,651 gal. each, 7 day residence 

total, 90% wv, API, atmospheric $ 57,488,911

Brush Cathode 0.08 mm SS316 Wire $ 400,686,098

Brush Anode 0.0072 mm Carbon Threads $ 62,316,288

Power Supply Based on Electrolyzer Power Supply, 

Scaled for 122,184 kW $ 18,550,000

MEC Transfer Pumps CS. 41000 gpm, 70 ft head. $ 2,906,592

TOTAL $541,705,673

22.3.3 Storage Subassembly

The storage subassembly for the MEC system is substantially smaller than for the fermentation 

system. The only storage needed is for the acetate feedstock and process water. The

components of this subassembly in general terms are listed in Figure 22-11 and shown in Figure 

22-12. 

Figure 22-11. Storage Components

Storage Subassembly Components

Acetate Storage Tank

Acetate Transfer Pump

Process Water Circulating Pump

Process Water Tank

Process Water Transfer Pump
154 

Figure 22-12. Storage Subassembly Design

Acetate

Storage Tank

Process Water

Tank

To MEC

To MEC

Acetate

Transfer Pump

Process Water

Transfer Pump

The equipment in Figure 22-12 is all very commonly available within the agricultural industry. 

Thus pricing for the equipment was easily located in reference documents and through 

discussions with suppliers. The capital cost of the subassembly is $884,167 as shown in 

Figure 22-13. 

Figure 22-13. Capital Costs of Storage Subassembly 

Storage Components Material Chosen Total Pricing

Acetate Storage Tank

SS316. 18697 gal, 90% wv, 12' dia x 22' 

high, atmospheric

 

$ 35,270 

Acetate Transfer Pump SS316. 215 gpm, 150 ft head $ 391,000 

Process Water Circulating Pump CS. 1199 gpm ea, 75 ft. head $ 161,593 

Process Water Tank CS. 756000 gal. 8 hr res time $ 54,088 

Process Water Transfer Pump CS. 41000 gpm, 70 ft head. $ 242,216

TOTAL $884,167

22.3.4 Gas Compression and Separation Subassembly

The function of gas compression and separation subassembly of the MEC is to compress the 

gaseous outputs from the MEC, to separate the product hydrogen, and to deliver it to the 

production facility limits at 300 psi. The components of this subassembly in general terms are 

listed in Figure 22-14. 
155 

Figure 22-14. Gas Compression and Separation Components

Gas Capture Subassembly Components

MEC Compressor (2 stage)

MEC PSA

MEC Compressor Intercooler 1

MEC Compressor Intercooler 2

Hydrogen Flow Meter

The outlet pressure of hydrogen at the plant gate of 300 psi was selected to provide with a 

system comparable to other DOE H2A Production Plants. 

At the MEC hydrogen outlet, the gas is at atmospheric pressure and 30oC. Since the 

temperature is only 30oC, there is a substantially lower mole fraction of water than is in the 

outlet gas from the fermentor, so no condenser is needed. In addition, the MEC electrolysis 

process separates most of the hydrogen from the CO2. It is estimated that only 5% of the MECproduced CO2 is mixed in the output hydrogen stream. The MEC-exit stream into the 

compressor is:

• H2 92,721 kg/day 

• CO2 50,607 kg/day 

• H2O 36,752 kg/day 

The gas mixture is sent to the 2-stage compressor. After the compressor and two intercoolers, 

the remaining gas mixture at 302 psi enters the PSA. Pressure loss in the PSA and the pipeline 

is estimated as 2 psi. A gas/water intercooler (MIC-1) is used between first and second stage 

compression to reduce the gas temperature to 45o

C and to condense out water. A second 

intercooler (MIC-2) is used between second stage compression and the PSA to reduce the gas 

temperature to 45o

C and to condense additional water, so that a total of 89% of the water vapor 

from the MEC output is removed before the MEC PSA., with the remaining water vapor mole 

fraction reduced to 0.00460. The gas/water intercoolers were scaled from a comparable heat 

exchanger used in the Ethanol Report. 

Since the PSA input is 1,914 moles/hr H2, 48 moles/hr CO2, and 9 moles/hr H2O, we have 

assumed that the PSA can recover 95% of the inlet H2. Thus the PSA output is: 

• H2 (high purity) 88,085 kg/day 

• Waste:

o CO2 50,607 kg/day 

o H2O 3,921 kg/day 

o H2 (loss) 4,636 kg/day 

Subassembly components are shown in Figure 22-15. 
156 

Figure 22-15. Gas Compression and Separation Subassembly Design

 

Hydrogen

Pipeline

Hydrogen

Flow Meter PSA

2 Stage

Compressor

Intercoolers

Gas compression and separation is has been priced previously in several reports. Thus pricing 

for the equipment was easily located in reference documents and through discussions with 

suppliers. The capital cost of the subassembly components is $16,099,326 as shown in 

Figure 22-16. 

Figure 22-16. Capital Cost of Gas Compression and Separation Subassembly 

Gas Compression and 

Separation Components

Material Chosen Pricing

Compressor Source: Using H2A Cost guidelines 

and scaling factors

$15,267,690

PSA Pressure Swing Adsorption $182,960

Source: Using H2A Cost guidelines and scaling factors

MEC Intercooler 1 Shell-Tube, SS304, 4 atmospheres $ 88,542

MEC Intercooler 2 Shell-Tube, SS304, 20 atmospheres $ 554,634

Hydrogen Flow Meter Information from Emerson Process 

Management $5,500

Total $16,099,326

22.4Bill of Materials

The flow chart outlining mass flows and subsystem parameters is shown in Figure 22-17. The 

detailed mass and energy balance is included in Part III, Appendix B. 
157 

Figure 22-17. MEC System Flow Chart 

 

The full bill of materials for the MEC stand-alone hydrogen production plant is shown in

Figure 22-18. The MEC system described in this section of the report has a capital cost of 

$558,689,165. 

Figure 22-18. MEC Stand-Alone Bill of Materials

 

Description Install Factor Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost

MEC Subassembly

MEC Tanks 1.3 $ 0.51 1.13E+08 $ 57,488,911

Brush Cathode 1.2 $ 14.60 27436736 $ 400,686,098

Brush Anode 1.2 $ 7.92 7868218 $ 62,316,288

Power Supply 1.3 $ 350,000.00 53 $ 18,550,000

MEC Pumps 2.8 242,216 11 $ 2,664,376

Storage Subassembly

Process Water Circulating Pump 2.8 17,635 2 $ 35,270

Process Water Tank 1.4 195,500 2 $ 391,000

Acetate Product Storage Tank 1.4 $ 161,593 1 $ 161,593

Acetate Transfer Pump 1.3 $ 54,088 1 $ 54,088

Process Water Transfer Pump 2.8 242,216 1 $ 242,216

Gas Capture Subassembly

Compressor MEC 1.3 1654 kgmol/hr 1654 kgmol/hr $ 9,233 1 $ 15,267,690

MIC - 1 1.3 $ 88,542 1 $ 88,542

MIC - 2 1.3 $ 554,634 1 $ 554,634

PSA MEC 1.3 1 $ 182,960

Hydrogen Flow Meter 1.3 1 $ 5,500 1 $ 5,500.00

System Initial Cost $ 558,689,165

MEC System - Batch Process

Efficiencies- MEC: 90% 

Acetate 

Storage

acetate

MEC

346,682,121 kg/day

30C CO2 Vented

1 atm

H2: 92,721 kg/day

CO2: 50,607 kg/day

VH20 36,756 kg/day

MEC

Compressor Intercooler liquid

stage 1 MIC-1 water

60 psi 180,084 kg/day 16,076 kg/day

180,084 kg/day 45C, 60 psi

stage 2 H2: 88,085 kg/day

302 psi H2 PSA

164,007 kg/day Intercooler CO2 147,249 kg/day

MIC-2 VH2O

164,007 kg/day PSA Waste Gas

45C liquid CO2: 50,607 kg/day

302 psi water H2O: 3,921 kg/day

16,758 kg/day H2: 4,636 kg/day

767,277 kg/day

water

Separation

345,540,502 kg/day and Recovery

unreacted liquids 

Water

Supply

345,914,844 kg/day 961,535 kg/day
158 

23.

In addition to the capital costs which have been described in the previous sections, there are 

some additional costs associated with this system, mostly electricity, labor, and various 

feedstock costs. These costs for both the MEC system and the Fermentation system are 

explained in detail in this section. 

Cost Assumptions and Calculations

23.1 Variable Costs

Aside from the fixed capital costs, there are numerous variable costs including electricity, labor, 

water usage, and additional consumables. The usage and associated costs of these will each be 

defined in the following sections. 

23.1.1 Electric Power Requirements

There are numerous elements of the Lignocellulosic Fermentation System requiring power, but 

they mostly center around pumps, conveyors and agitators. However, the largest power needs 

comes from the compressor, which is compressing an enormous amount of gas. PSA power is 

minimal as it is only needed to actuate valves. The gas compressor is a 2-stage piston 

compressor (N=2) with interstage cooling. Its power was calculated by assuming compression 

from 45o

C inlet temperature with efficiency of 75%. Overall pressure ratio was 20.5 for an 

outlet pressure of 302psi. 

Figure 23-1 lists the electricity usage anticipated for each of the Fermentor components

previously described and provides a total consumption value for the Fermentor to be used in 

further analysis. The previously mentioned duty cycles and operational cycle have been taken 

into account in these computations. We have assumed a pump efficiency of 80% and a density 

of all liquids as 1 g/cc. H2A accounts for electricity usage and costs in the “Other Variable 

Costs” section.
159 

Figure 23-1. Power Requirements for Fermentor without MEC

Conveyors Power Unit kW

Bale Transport Conveyor 50 HP 37

Bale Unwrapping Conveyor 50 HP 37

Belt Press Discharge Conveyor 50 HP 37

Shredder Feed Conveyor 50 HP 37

Lignin Wet Cake Screw 50 HP 37

DAP Solids Feeder 50 HP 37

Agitators Power Unit

Fermentation Agitator 75 HP 56

Hydrozylate Mix Tank Agitator 50 HP 37

Seed Hold Tank Agitator 23 HP 17

4th Seed Agitator 6 HP 4

5th Seed Agitator 19 HP 14

CSL Storage tank Agitator 23 HP 17

CSL/DAP Day Tank Agitator 5 HP 4

Pumps Throughput Unit Head M3/hr Meter

Wash Table Pump 2500 GPM 50 ft 568 15 29

Wash Water Pump 5000 GPM 50 ft 1136 15 59

Clarifier Underflow Pump 100 GPM 50 ft 23 15 1

Clarified Water Pump 5000 GPM 50 ft 1136 15 59

Belt Press Sump Pump 100 GPM 40 ft 23 12 1

Sulfuric Acid Pump 4 GPM 245 ft 1 75 0.2

Process Tank Pump 737 GPM 200 ft 168 61 35

Fermentation Recirculation Pump 1060 GPM 150 ft 241 46 38

CSL Pump 431 GPM 150 ft 98 46 15

Sulfuric Acid Pump 215 GPM 150 ft 49 46 8

Cooling Water Pump 41000 GPM 70 ft 9318 21 677

Make-up Water Pump 1083 GPM 75 ft 246 23 19

Process Water Circulating Pump 1199 GPM 75 ft 273 23 21

CSL/DAP Pump 431 GPM 150 ft 98 46 15

Seed Hold Transfer Pump 172 GPM 150 ft 39 46 6

Seed Transfer Pump 1231 GPM 100 ft 280 30 29

Miscellaneous Throughput Unit

DAP Unloading Blower 253 HP 189

Compressor 4535

PSA 50

Lime Unloading Blower 253 HP 189

Total Power Usage (kW) 6,350

kWhr/day 152,389

kWhr/year 55,621,907

$/year $ 2,781,095

$/kg H2 $ 0.23

kWhr/kg H2 4.10

kwhr/kwhr H2 0.12

Power Requirements

The power requirement of the MEC system comes primarily from the MEC cell itself. The 

power draw of the MEC electrolysis process is estimated from the electron current required to 

reduce the H+ and the applied voltage and the system electric efficiencies experimentally 
160 

determined in Penn State tests98. Figure 23-2 lists the electricity usage anticipated for the MEC 

system. 

Figure 23-2. Power Requirements MEC System

kW

MEC 122184

Compressor MEC 6619

PSA 50

Pumps Throughput Unit Head M3/hr Meter kW

Cooling Water Pump 41000 GPM 70 ft 9318 21 8125

Acetate Transfer Pump 1083 GPM 75 ft 246 23 19

Process Water Circulating Pump 1199 GPM 75 ft 273 23 42

kW 137,040

kWhr/day 3,288,962

kWhr/year 1,200,471,198

$/year $ 60,023,560

$/kg H2 $ 2.07

kWh/kg H2 37.34

kWh/kWh H2 1.12

Power Requirements

23.1.2 Labor

Labor rates for the fermentation system are similar to those in the ethanol report. We 

eliminated lab technicians and managers, since we will not require use of a lab. We also 

removed clerks and secretaries, who are accounted for separately in the H2A spreadsheet. Due 

to the vastly smaller and less complex plant as compared to an ethanol plant, the resulting labor 

numbers are probably somewhat high. However, given the relatively small impact that labor

has on the cost of produced hydrogen, we felt that this was an acceptable overestimation. 

Figure 23-3 shows labor categories and number of personnel. The same number of personnel is 

assumed for the fermentation-only alone plant as is assumed for the fermentor with MEC plant. 

H2A accounts for labor costs in the fixed O&M. 

Figure 23-3. Labor Breakdown - Fermentor

Position Number of Personnel

Plant Manager 1

Plant Engineer 1

Maintenance Supervisor 1

Shift Supervisor 5

Maintenance Technician 8

Shift Operators 20

Yard Employees 32

General Manager 1

Total FTEs 68

 

98Call, Douglas et al. “High Surface Area Stainless Steel Brushes as Cathodes in Microbial Electrolysis Cells.” 

Environmental Scientific Technology (2008) table 1.
161 

For MEC labor, due to the similar size but reduced complexity, we have assumed 1/4 of the 

labor is required.

23.1.3 Additional Chemical Consumables

There are many consumables that are needed for full operation of the lignocellulosic 

fermentation plant. The rates and costs of these consumables for the fermentation system are 

listed in Figure 23-4. Rates and costs for these consumables are derived from the ethanol 

production report. 

Figure 23-4. Additional Consumables Costs – Fermentation system

Raw Material Unit Rate Cost ($/kg) $/year $/kg H2

CSL kg/day 31,344 $ 0.18 $ 2,023,606 $ 0.165678

Sulfuric Acid kg/day 78,912 $ 0.03 $ 785,743 $ 0.064331

Lime kg/day 57,480 $ 0.08 $ 1,606,244 $ 0.131508

DAP kg/day 3,912 $ 0.16 $ 221,778 $ 0.018158

Propane kg/day 480 $ 0.005 $ 848 $ 0.000069

Clarifier Polymer kg/day 672 $ 2.75 $ 674,520 $ 0.055225

Totals $ 5,312,739 $ 0.43

Additional Consumables Cost

Given that these consumables are not options within the H2A Modeling tool, we have 

calculated the cost per kilogram of hydrogen from these variable operating costs and manually 

added these to the results obtained from H2A for levelized hydrogen costs. 

The MEC consumable for the stand-alone system is acetic acid at $0.595/kg and costing $5.70 

per kg of produced hydrogen. 

23.1.4 Water Consumption

There are several sources of water loss within each system. For the fermentation system, water 

is primarily lost in the fermentation reaction. Further water is trapped in solids that are removed 

as waste. Most of the process water vapor is returned by condensation in the condenser and 

intercoolers. Using the molar ratios of the fermentation equation and the volume of hydrogen 

gas generated, we are able to calculate the daily use of reactant water in hydrogen production. 

In addition, substantial water is recovered from liquid byproducts. 

For the MEC developed by PSU, the substrate is much more dilute than the fermentor output. 

Further, additional water is needed because we have assumed that after going through ten cycles 

of MEC acetate replenishment, the system will be flushed and the process water will need to be 

replaced. Thus, every day, 10% of the overall process water will need to be replaced in addition 

to the water used up in the production of hydrogen. 

The water losses for the Fermentor and MEC is shown in Figure 23-5. 
162 

Figure 23-5. Water Use

Water Loss (gal/day) 

Fermentor MEC

Size 37 TPD 88 TPD

Water Used in reaction 54,826 109,375

Water Loss in Lignin 

Separator 223,397 NA

MEC Waste Water (10%) NA 9,132,152

Water Lost in PSA 759 1,035

Net Loss 278,983 9,241,551 

Net Loss (gal/kgH2) 7.5 105

23.1.5 Total Variable Feed Costs

The sum of all the variable costs, excluding labor, for the Fermentation system is indicated in 

the Figure 23-6. The variable costs for the MEC system are shown in Figure 23-7. 

Figure 23-6. Total Variable Costs for Fermentor 

Raw Material Unit Rate Cost ($/kg) $/year $/kg H2

Corn Stover kg/day 2,000,000 $ 0.03 $ 24,090,000 $ 1.972317

Process Water gal/day 278,983 $ 0.004 $ 373,089 $ 0.030546

CSL kg/day 31,344 $ 0.18 $ 2,023,606 $ 0.165678

Sulfuric Acid kg/day 78,912 $ 0.03 $ 785,743 $ 0.064331

Lime kg/day 57,480 $ 0.08 $ 1,606,244 $ 0.131508

DAP kg/day 3,912 $ 0.16 $ 221,778 $ 0.018158

Propane kg/day 480 $ 0.005 $ 848 $ 0.000069

Clarifier Polymer kg/day 672 $ 2.75 $ 674,520 $ 0.055225

Totals $ 29,775,828 $ 2.44

Variable Operating Costs

Figure 23-7. Total Variable Costs for MEC 

Raw Material Unit Rate Cost ($/kg) $/year $/kg H2

Acetic Acid kg/day 767,277 $ 0.595 $ 166,633,382 $ 5.758714

Process Water gal/day 9,242,551 $ 0.002 $ 5,618,279 $ 0.194163

Totals $ 172,251,661 $ 5.95

Variable Operating Costs
163 

24.

The summation of the various costs per kg of hydrogen produced is listed in 

Hydrogen Production Costs for Fermentor and MEC

Figure 24-1. Note 

that due to slight differences in H2A costing methodology, the prices shown in these subtotals 

might be different from those shown in the final results. The primary components of the 

levelized cost of hydrogen are capital costs, fixed O&M and other variable costs as shown in 

Figure 24-2. 

Figure 24-1: Total Hydrogen Cost 

Cost Component

System Fermentor-Only 

(37 TPD Production)

MEC Alone (88 TPD)

(at 0.9V/cell)

Cost Component Hydrogen Production Cost 

Contribution ($/kg)

Hydrogen Production Cost 

Contribution ($/kg)

Capital Costs $1.19 $4.37

Decommissioning Costs $0.01 $0.05

Fixed O&M $0.87 $0.71

Feedstock Costs $1.60 $5.18

Other Raw Material Costs $0.00 $0.00

Byproduct Credits $0.00 $0.00

Other Variable Costs (including 

utilities) $0.23 $2.12

Additional Consumables $0.43 $0.00

Total $4.33 $12.43

Hydrogen Production Cost Contribution ($/kg)

Total Cost of Produced H 2

Figure 24-2. Comparison of Levelized Cost Components 

$0.00

$2.00

$4.00

$6.00

$8.00

$10.00

$12.00

$14.00

Fermentor-Only 

(37 TPD Production)

MEC Alone (88 TPD)

(at 0.9V/cell)

Additional Consumables

Other Variable Costs 

(including utilities)

Feedstock Costs

Fixed O&M

Decommissioning Costs

Capital Costs

For the Stand-alone Fermentor system, there is a substantial liquid organic byproduct which has 

value, but for which we have not taken a cost credit due to the uncertainty of the constituents 
164 

and the market values. From NREL testing, the fermentor byproduct (section 2.3.2) is as shown 

in Figure 24-3. 

Figure 24-3. Fermentor Liquid-Organic Output Value

Fermentor Product mM MW Mass fraction Market Cost per kg

Acetic Acid 26.0 60.05 51.1% $0.57

Ethanol 14.0 46.06 21.1% $0.54

Succinic Acid 5.6 118.10 21.6% NA

Lactic Acid 1.8 90.08 5.3% $2.03

Formic Acid 0.6 43.03 0.8% NA

Thus, after a distillation separation process, the resulting components would have a current 

average market value of about $0.55/kg. However, any large addition to the world market

would reduce prices, so the actual cost benefit of the organic byproduct is difficult to estimate. 

With the fermentor organic liquid byproduct of about 700,000 kg/day, if the value of this raw 

byproduct were $0.20/kg, the byproduct value would be $140,000 per day. This would yield a 

credit of $3.78/kg H2, substantially driving down the H2 cost/kg to $0.60/kg. Even with a 

byproduct value as low as $0.12/kg, the H2 cost/kg is $2.09/kg, near the goal of $2.00/kg H2. 

25.

25.1 Lignocellulose Fermentation System

Lignocellulose Fermenation and MEC System Conclusions and 

Recommendations 

The lignocellulose fermentation system used corn stover at 2,000 tonnes/day as a feedstock to 

generate 37 tonne/day purified H2. The process is complex and involves feedstock preprocessing, hydrolysis, saccharification, fermentation, solid/liquid separation, and H2 gas 

separation. The unique 150oC acid hydrolysis process used to convert cellulose and 

hemicellulose into complex sugars has been demonstrated by NREL. The subsequent 

saccharification process utilizes organisms developed at NREL to rapidly convert the hydrolysis 

products into simple sugars. The ensuing fermentation process occurs in the same reactor and 

uses separate NREL-developed bacteria to produce H2 and CO2 gases plus simple organic 

liquids, such as acetic acid and ethanol. 

In the initial calculation, with no value recovered from the byproducts, the lignocellulosic 

processing achieved a moderately low H2 cost of $4.33/kg H2. For this system producing 37 

tonnes of pure H2, these byproducts are essentially: 692 tonnes of organic acid and alcohol 

liquids, 115 tonnes of unreacted sugars, and 846 tonnes of solids including lignin and a small 

amount of unreacted xylan. The liquid organic byproducts, equaling 35% of the input organic 

feedstock by weight, include 51% acetate and 21% ethanol which have intrinsic value, but 

require a subsequent separation process such as distillation. As an alternative to byproduct 

recovery, the liquid organics can be processed in an MEC to produce additional hydrogen. 

There is high potential for H2 cost reduction from sales of the liquid organic byproduct. If this 

byproduct had a market value of $0.12/kg (as compared to the acetic acid market price of ~ 

$0.60/kg), the net H2 cost would be reduced to $2.09/kg . 
165 

In addition to the liquid byproducts, the remaining solids, lignin and some residual xylan, (42% 

of the organic feedstock by weight) have additional value, as they can be used as a low grade 

fuel for industrial processes. 

Recommendations: 

• Conduct further analysis to verify the values of the lignocellulose fermentation liquid 

and solid organic byproducts to significantly reduce the cost of the hydrogen. 

• Conduct further analysis to address the separation of byproduct components (acids, 

ethanol, etc.), such as in a distillation process

• Examine in greater detail the current experimental efficiencies and reaction times for the 

hydrolysis, saccharification, and fermentation processes, and propose potential future 

testing to verify these quantities.

25.2 Microbial Electrolysis Cell (MEC) System

For the MEC electrolysis process the microbial electrolysis voltage is supplemented by an 

external voltage. Low external voltage results in highest efficiency, but increases reaction 

times, thus increasing system size and cost. For lowest H2 production cost, the system should 

be operated at a higher than normal voltage (0.9V) to minimize capital plant costs. The very 

low concentration of the input acetic acid ~2 g acetate/liter of water, (as opposed to the 

fermentor at 200 g organics/liter of water) resulted in a very large system, i.e., 96,000,000 

gallons of total reactor volume for 88 tonnes pure H2/day. These volumes, based on current 

PSU research results, resulted in large tankage and the resultant large area anodes and cathodes, 

which were the major cost drivers. In addition, the effective operation of these electrodes in a 

very large scale reactor tank has not yet been analyzed or demonstrated. 

For the MEC system, the H2 cost using pure acetic acid feedstock was a moderately high 

$12.43/ kg H2. This was primarily a product of the use of a very dilute acetic acid/water 

reactant, which necessitated very large reactor volumes and very large, costly anodes and 

cathodes. Along with the high capital cost was high acetate feedstock cost, which could be 

reduced by using a low purity simple organic feedstock such as the fermentor byproduct. 

However, the current large system size dominates the costs. The immaturity of the full scale 

system concepts and components indicated that there is extensive potential for future cost 

reductions. 

Recommendations: 

• Develop optimized production process and components for low capital cost systems, 

addressing issues such as:

o Increased solution density, 

o Pressurized operation (to minimize separate compressor costs)

o Lower cost cathodes 

o Lower cost anodes

o Cathode and anode geometry optimization for large reactors 

• Determine the extent of ion transport losses as reactor size grows to production scale

• Examine potential efficiencies of the process carried out at large scale
166 

26.

Process Mass flow mole flow T T ave ≅T P Cp h dq/dt

% kg/day kg/hr kgM/hr °C °C °C psi kJ/kg °C kJ/kg kJ/hr

Feedstock Corn stover 20.0% 2,000,000 83,333 20 14.7

Inputs cellulose 36.7% 734,000 30,583

hemicell 27.4% 548,000 22,833

Lignin,etc 35.9% 718,000 29,917

total 100.0% 2,000,000 83,333

water 80.0% 8,000,000 333,333 20 14.7

total slurry 100.0% 10,000,000 416,667 20 14.7

HX-1 from feedstock 10,000,000 416,667 20.0 14.7 4.180

to HX-2 10,000,000 416,667 43.3 31.7 23.3 40,655,319

from fermentor 9,332,641 388,860.06 55.0 14.7 4.182

to LigSep 9,332,641 388,860 30.0 42.5 -25.0 -40,655,319

HX-2 from HX-1 primary 10,000,000 416,667 43.3 68.8 4.206

to hydrolyzer heater 10,000,000 416,667 128.5 85.9 85.2 149,281,250

from hydrolyzer 10,000,000 416,667 140.0 85.0 68.8 4.215

to sacchrification 10,000,000 416,667 55.0 97.5 -149,281,250

Hydolyzer heating of slurry from HX-2 10,000,000 416,667 128.5

Heater 10,000,000 416,667 150.0 139.3 21.5 68.8 4.290 38,387,188

heating of CO2 from PSA 490,873 20,453 45.0

CO2/H2mass 54.58 490,873 20,453 150.0 97.5 105.0 14.7 0.913 1,960,729

heating of H2O from PSA 2,783 116 45.0

H2O/H2mass 0.31 2,783 116 150.0 97.5 105.0 14.7 1.970 23,985

Total heat input required (neglecting H2 ≅T) 40,371,902

Waste H2 burner heating effic =

H2 LHV kJ/kg = 119,705 90% 8,994 375 14.7 40,371,902

 10,000,000 416,667

Total heating from HX-1 + HX-2 + Heater 230,308,471

Hydrolyzer In from Heater 10,000,000 416,667 150.0 68.8

inputs cellulose 36.7% 734,000 30,583 psi

hemicell 27.4% 548,000 22,833 4.68

Lignin,etc 35.9% 718,000 29,917 atm

water 80.0% 8,000,000 333,333

total 10,000,000 416,667

efficiency= output complex sugars 1,153,800 48,075 140.0 68.8

90% xylan 128,200 5,342

Lignin,etc 718,000 29,917

organics 2,000,000 83,333

water 8,000,000 333,333

total 10,000,000 416,667

Endothermic heat absorption ≅T 10.0

to HX-2 10,000,000 416,667 140.0 68.8

Saccharification from HX-2 10,000,000 416,667 55.0 14.7

in fermentor tank to fermentation process 10,000,000 416,667 55.0

Fermentor in from total 10,000,000 416,667 55.0 14.7

efficiency = sacchrification simple sugars 1,153,800 48,075 267

(glucose)

90% process Lignin,xylan,etc 846,200 35,258

water 8,000,000 333,333

reactants 10,000,000 416,667

H2O reacted in fermentation 207,676 8,653

gas out H2 46,476.6 1,937 961 55 14.7

CO2 507,339.5 21,139 480 55

V H2O 113,542 4,731 263 55

total gas 667,359 27,807 1,704

slurry out to HX-1 slurry total 9,332,641 388,860 55

Lignin,xylan,etc 846,200 35,258

Acetate 692,280 28,845 

Glucose 10% 115,380 4,808 

L-H20 7,678,781 319,949

total 9,332,641 388,860

total mass out 10,000,000 416,667

Condenser gases in H2 46,477 1,937 55 14.7 14.496

FC-1 from fermentor CO2 507,340 21,139 55 14.7 0.953

V H2O 113,542 4,731 55 14.7 1.912

total 667,359 27,807

gases out to H2 46,477 1,937 960.6 45 50 -10 14.7 14.496 -280,719

Compressor-F-1 CO2 507,340 21,139 480.3 45 50 -10 14.7 0.953 -201,350

V H2O 64,477 2,686.5 149.1 45 50 -10 14.7 1.912 -51,367

total gas 618,293 25,762 -533,436

liquids out L H2O 49,066 2,044 45 -10 14.7 4.200 -62,477

V-L Phase change heat -2222 -4,542,656

Part III Appendix A: Fermentor Mass/Heat Balance
167 

Process Mass flow mole flow T T ave , T P Cp h dq/dt

% kg/day kg/hr kgM/hr °C °C °C psi kJ/kg °C kJ/kg kJ/hr

Compressor in H2 46,477 1,937 961 45 14.7

stage 1 CO2 507,340 21,139 480 45 14.7

F-1 V H2O 64,477 2,687 149 45 14.7

total 618,293 25,762 1,590 45 14.7

out H2 46,477 1,937 961 190.8 60

CO2 507,340 21,139 480 190.8 60

V H2O 64,477 2,687 149 190.8 60

total 618,293 25,762 1,590

Intercooler Gas in H2 46,477 1,937 961 190.8 60 14.496

FIC-1 CO2 507,340 21,139 480 190.8 60 0.953

 Primary V H2O 64,477 2,687 149 190.8 60 1.912

total 618,293 25,762

Gas out H2 46,477 1,937 960.6 45 117.9 -146 58 14.496 -4,092,881

CO2 507,340 21,139 480.3 45 117.9 -146 58 0.953 -2,935,688

V H2O 15,169 632.1 35.1 45 117.9 -146 58 1.912 -176,199

total gas 568,986 23,708 -7,204,769

Liquid out L H2O 49,307 2,054 45 -146 58 4.200 -915,401

V-L Phase change heat -2222 -4,565,040

total out 1,187,279 49,470 -12,685,209

 Secondary water in 3,642,123 151,755 20 4.180

water out 3,642,123 151,755 40 30 20 4.180 12,685,209

Compressor in H2 46,477 1,937 45 58

stage 2 CO2 507,340 21,139 45 58

F-2 V H2O 15,169 632 45 58

total 568,986 23,708 1,476

out H2 46,477 1,937 960.6 218 302

CO2 507,340 21,139.15 480.3 218 302

V H2O 15,169 632 35.1 218 302

total 568,986 23,708 1,476.0

FIC-2 Gas in H2 46,477 1,937 218 302

 Primary CO2 507,340 21,139 218 302

V H2O 15,169 632 218 302

total 568,986 23,708

Gas out H2 46,477 1,937 960.6 45 131.45 -173 300 14.496 -4,853,629

CO2 507,340 21,139.15 480.3 45 131.45 -173 300 0.953 -3,481,348

V H2O 2,876 120 6.7 45 131.45 -173 300 1.912 -39,619

total gas 556,692 23,196 1,447.6 -8,374,596

Liquid out L H2O 12,293 512 45 -173 300 4.200 -270,646

V-L Phase change heat -2222 -1,138,145

total out 568,986 23,708 -9,783,386

 Secondary water in 2,808,964 117,040 20 4.180

water out 2,808,964 117,040 40 30 20 4.180 9,783,386

PSA in total 556,692 23,196 45 300

 fermentor H2 46,477 1,937 960.6 45 300

CO2 507,340 21,139 480.4 45 300

V H2O 2,876 120 6.7 45 300

out H2 outlet 80% 37,181 1,549 45 300

H2 loss 20% 9,295 387 192.1 45 300

CO2 507,340 21,139 480.4 45 300

CO2/H2 mass ratio 54.6

V H2O 2,876 120 6.7 45 300

VH2O/H2 mass ratio 0.31

waste gas out 519,511 21,646 679.2

Burner use of waste H2= 97%

LigninSeparator in from HX-1 total from HX-1 9,332,641 388,860

Lignin,etc 846,200 35,258 30 14.7

Acetate 692,280 28,845

Glucose 115,380 4,808 

L-H20 7,678,781 319,949

total 9,332,641 388,860

out - mash Lignin,etc 100% 846,200 35,258 30

Acetate 5% 34,614 1,442

Glucose 5% 5,769 240

L-H2O = lignin wt 846,200 35,258

total 1,732,783 72,199

out - liquid Acetate 95% 657,666 27,403 30

to MEC Glucose 95% 109,611 4,567

H20-L 6,832,581 284,691

total 7,599,858 316,661

out total 9,332,641 388,860
168 

27.

 Mass flow mole flow T T ave ∆T P Cp h dq/dt

% kg/day kg/hr kgM/hr °C °C °C psi kJ/kg °C kJ/kg kJ/hr

MEC Inputs: Acetate 767,277 31,970 532.4 30 14.7

0.9V system Acetate water input 345,914,844 14,413,119

 dilution (g/L) = H2O + substrate = 346,682,121 14,445,088

2.213 total input 346,682,121 14,445,088

reactant vol.=

reacted acetate 90% 690,549 28,773 479.1

reacted water 414,314 17,263 958.3

90% output-gas H2 92,721 3,863 1,916.5 30 14.7

reaction CO2 total out 1,012,142 42,173 958.3 14.7

completion CO2 mixed in H2 5% 50,607 2,109 47.9 0.05

H2O-V 36,755 1,531.5 85.0 14.7

total cathode gas 1,141,619 47,567 2,049

output liquid unreacted acetate 10% 76,728 3,197 30

H20-L 345,463,775 14,394,324

total liquid 345,540,502 14,397,521

total output 346,682,121 14,445,088

CO2 Vented from MEC 95% 961,535 40,064

Compressor in H2 92,721 3,863 1,916 30 14.7

MC-1 CO2 50,607 2,109 48 30 14.7

V H2O 36,755 1,531.5 85 30 14.7

total 180,083 7,503 2,049 30 14.7

out H2 92,721 3,863 1,916 204 60

CO2 50,607 2,109 47.9 204 60

V H2O 36,755 1,531 85 204 60

total 180,083 7,503 2,049 60

Intercooler Gas in H2 92,721 3,863 204 60

MIC-1 CO2 50,607 2,109 204 60

 Primary V H2O 36,755 1,531 204 60

total 180,083 7,503

Gas out H2 92,721 3,863 1,916 45 124.3 -159 58 14.496 -8,882,141

CO2 50,607 2,109 47.9 45 124.3 -159 58 0.953 -318,543

V H2O 20,679 861.6 47.8 45 124.3 -159 58 1.912 -261,287

total 164,007 6,834 -9,461,971

Liquid out L H2O 16,076 670 45 -159 58 4.200 -324,656

V-L Phase change heat -2222 -1,488,369

total out 180,083 7,503 -11,274,996

 Secondary water in 3,237,228 134,885 20 4.180

water out 3,237,228 134,885 40 30 20 4.180 11,274,996

Compressor in H2 92,721 3,863 45 58

MC-2 CO2 50,607 2,109 45 58

V H2O 20,679 862 45 58

total 164,007 6,834 2,012 45 58

out H2 92,721 3,863 1,916 267 302

CO2 50,607 2,109 48 267 302

V H2O 20,679 862 48 267 302

total 164,007 6,834 2,012

Intercooler Gas in H2 92,721 3,863 267 302

MIC-2 CO2 50,607 2,109 267 302

 Primary V H2O 20,679 862 267 302

total 164,007 6,834

Gas out H2 92,721 3,863 1,916 45 156 -222 300 14.496 -12,432,757

CO2 50,607 2,109 48 45 156 -222 300 0.953 -445,880

V H2O 3,921 163.4 9.1 45 156 -222 300 1.912 -69,347

total gas 147,249 6,135 -12,947,984

Liquid out L H2O 16,758 698 45 -222 300 4.200 -473,726

V-L Phase change heat -2222 -1,551,548

total out 164,007 6,834 -14,973,258

 Secondary water in 4,299,057 179,127 20 4.180

water out 4,299,057 179,127 40 30 20 4.180 14,973,258

PSA in H2 92,721 3,863 1,914 45 300

 MEC CO2 50,607 2,109 48 45 300

V H2O 3,921 163 9 45 300

total 147,249 6,135 45 300

out H2 out 95% 88,085 3,670 45 300

H2 waste 5% 4,636 193 45 300

CO2 50,607 2,109 48 45 300

V H2O 3,921 163 9 45 300

total out 147,249 6,135

91,524,080 gallons

Part III Appendix B: MEC Mass/Heat Balance
169 

Part IV: Integrated Systems

 
170 

28.

Having characterized and assessed costs for several pathways for biological hydrogen 

production, it is important to consider how combinations of different pathways could enhance 

hydrogen production and potentially lower costs. Combining capital costs, making use of waste 

products and increasing solar capture makes it possible to improve the cost-effectiveness of the 

overall system. This report is focused on examining different combinations of the biological 

hydrogen production pathways and evaluating costs associated with these integrations as 

compared to the individual systems. 

Systems Integration

There are three types of biological systems that have been evaluated in this effort: 

photobiological, fermentation, and Microbial Electrolysis. These types can be used in 

integrated systems. Integrations are shown in Figure 28-1. Combinations can occur within the 

photobiological subset of systems, within the fermentation subset of systems or a combination 

of a photobiological and fermentation system. While other combinations may exist, our task 

required the evaluation of four systems. We have chosen to study the four that are most 

representative and have sufficient synergies to make integration a possibility. Higher multiples 

of the systems could be possible but are not considered necessary to evaluate in the boundary 

analysis. Those could be reviewed in more detail if any of these combinations prove fruitful. 

Figure 28-1. Potential Biological Integration Methods

B-3 Algal Sulfate 

PermeaseMutant

[top Reactor Bed] 

B-5 PNS Bacteria

[lower reactor bed 

Lignocellulosic

Fermentation

MEC

B-3 Algal Sulfate 

Permease

Reactor Bed 

Algal 

Fermentation 

(C-3)

B-5 PNS

Reactor Bed 

Algal 

Fermentation 

(C-5)

Raw 

acetate 

byproduct

Acetate

(~5% of 

req’mt)

Waste

organism

Waste

organism

Acetate

The four potential system combinations analyzed are as follows: 

• stacked photobiological systems (B-3/B-5) to capture a greater portion of the solar 

spectrum

• photobiological H2 production (B-1) combined with algal fermentation (C-1) 

• photobiological PNS H2 production (B-5) combined with waste PNS fermentation (C-5) 

and waste acetate consumption in the PNS photosynthesis 

• lignocellolosic fermentation combined with microbial electrolysis

For more details on the specifics of each of the systems mentioned please refer to Part I, Part II 

and Part III of this report. 
171 

28.1 Stacked Photobiological Pathway

The best way to evaluate integrated photobiological beds is by considering them vertically 

stacked. Side-by-side integration would increase the land requirements and only share auxiliary 

subsystem costs. Organisms and bed designs as defined would not absorb the full spectrum or 

intensity of solar light available. Thus, it seems probable that a stacked system would take 

advantage of this and allow for greater hydrogen production per area. In order to stack systems 

the organisms chosen for the top and bottom reactor must have sufficiently different PAR 

spectra to produce significant increases in hydrogen production in a stacked configuration. As 

has been noted before, the PNS bacteria in the B-5 pathway utilize a different PAR spectrum 

than the Chlamydomonas algae and Cyanobacteria in systems B-1, B-2, B-3, and B-4. Thus 

stacking one of these beds on top of a B-5 photobio bed could improve H2 production by 

increasing the range of the solar spectrum captured by the beds. For the upper bed, B-3 was 

chosen rather than the B-1 and B-2 systems because of the additional greater precautions 

necessary in handling the stoichiometric H2/O2 mix from the latter systems. The B-4 system 

was not used because the polypropylene mats used for the immobilized algae, combined with

the high algae density on these mats, would block the transmission of photons to the lower 

layer. Therefore, the stacked system we chose to investigate consisted of a sulfate permease 

mutant B-3 system stacked on top of a B-5 PNS bed, increasing the effective PAR of the 

integrated system from 44% to 71%, leading to an increase in hydrogen production relative to 

the B-3 system by itself. Furthermore, many hardware components can be combined for the 

two systems, including most of the reactor bed subassembly, control system subassembly, and 

gas capture subassembly costs. The assumptions and changes from the B-3 system to create a 

stacked bed include:

 

• A second layer of film to separate the two beds 

• The components of the Organism Feed Subassembly and Recycle Subassembly are all 

doubled in order to provide nutrients and maintain concentration in both beds 

• In order to capture the additional photons, roughly 1/5 of the normal concentration of 

PNS will be needed, and thus the acetate nutrient requirements are 1/5 of the full B-5 

system

• Since the B-5 system will be stacked below the B-3 system, it will be impossible to 

provide paddlewheel mixing. Thus, we have added recirculation pumps and perforated 

pipe that extends the length of the raceways every 10 feet 

A process diagram of the integrated system is shown in 

Figure 28-2. The conceptual design of a tiered reactor bed system is illustrated in Figure 28-3. 

Overall cost for the B-3/B-5 stacked system for a 1 ton per day (TPD) plant is $4,274,085. The 

bill of materials for the B-3/B-5 stacked system is shown in Figure 28-4. 
172 

Figure 28-2. Process Diagram – Integrated Photobiological Systems

Photo Bio Reactor Bed

Mixers

Power

Source

LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Control Room PC

Instrumentation

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Raceway

Wiring Panel

Instrumentation

(x 20 Raceways)

Paddlewheel Motors

(x 20 Raceways)

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Transfer Out

Pump

Wet Spent

Organisms

Makeup

Water

Algal

Separation

Recycle

Subassembly

CO2

Cylinder

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve

Transfer Out

Pump

Makeup

Water

Recycle

Subassembly

CO2

Cylinder

Nutrient /

Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve

Organism Feed

Subassembly

Figure 28-3. Tiered Reactor Bed Conceptual Design

Transparent Film

Earthen Berms Raceways

Gas Tight Frame

Film Attachment Points Driveway

Earth Trough Bed Depth

Paddlewheel

Inlet Ports: Makeup

Water & Nutrients

FRONT VIEW

Perforated

Mixing Pipes
173 

Figure 28-4. Bill of Materials for B-3/B-5 Stacked Beds Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 165452 m2 1 m2 $ 0.54 165,452 $ 89,054 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Film Layer 2 165452 m2 1 m2 $ 0.54 165,452 $ 89,054 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 178597 m2 1 m2 $ 0.47 178,597 $ 84,382 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 

Pond Edging 26566 m 1 m $ 7.00 26,566 $ 185,965 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 38 $ 991,154 Based on California labor rates.

Paddlewheel Mixers 1 each $ 5,000.00 76 $ 380,000 Est. paddlewheel cost. 4 per raceway. 18 raceways.

Inlet Water Valve 0.5 in 1 each $ 43.46 38 $ 1,651 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 38 $ 1,651 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 38 $ 2,555 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 114 $ 912 Price from http://www.ancorp.com/line.aspx?id=819

PNS Recicirculaiton Pump 76 $ 187.99 76 $ 14,287

PNS Perforated Piping 248900 $ 0.24 248,900 $ 59,736

Gas Capture Subassembly

Compressor 67 kgmol/hr 67 kgmol/hr $ 9,233.00 1 622,078 $ Using H2A Unit cost

PSA $ 78,900 1 78,900 $ 

Raceway Collection Pipe 76 ft 1 ft $ 1.00 76 $ 76 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, 

 Main Collection Pipe 2160 ft 1 ft $ 6.18 2160 13,349 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, 

 ( ) Gas Capture Pipe 350 ft 1 ft $ 6.18 350 2,163 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, 

 PNS Gas Collection Pipe $ 1.00 380 $ 380 Diverts PNS gas to Main Gas Piping

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 18 $ 184,536 Est. Based on Perry's Chemical Handbook, Plant Design and Economics

Nutrient Metering Pumps 1468 Gal/Hr. $ 2,594 2 $ 5,188 Est. Based on Perry's Chemical Handbook, Plant Design and Economics

Mixing Pump 150 gpm 150 gpm $ 10,252 18 $ 184,536 Est. Based on Perry's Chemical Handbook, Plant Design and Economics

CO2 Cylinder 50 lb 50 lb $ 360 1 $ 360

 y p p y 

(http://kegman.net/carbon.htm)

CO2 Saturation Vessel 15000 gal 15000 gal $ 30,000 10 $ 300,000

 y g 

Economics

Nutrient Tank 14 gal 10 gal $ 30.00 2 $ 60

 g y 

085 (http://www.watertanks.com/products/0275-085.asp)

Raceway Slurry Collection Pipe (B3 and B5) 184 ft 1 ft $ 2.12 184 $ 390 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and 

 / Main Slurry Collection Pipe (B3 and B5) 5280 ft 1 ft $ 8.51 5280 $ 44,933 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and 

 / Recycle-Feed Transfer Pipe (B3 and B5) 360 ft 1 ft $ 8.51 360 $ 3,064 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and 

 /

Recycle Subassembly

Rotary Drum filter 17980 gph 3750 gph $ 87,000.00 5 435,000 $ Information from Dana Kent at Advanced Aquaculture Inc. on Hydrotech Drum Filter

Transfer Out Pump 150 gpm 150 gpm $ 10,252 184,531 18 $ 

 y g 

Timmerhaus

Main Slurry Feed Pipe (B3 and B5) 5280 ft 1 ft $ 6.18 2040 12,607 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and 

 / Raceway Slurry Feed Pipe (B3 and B5) 184 ft 1 ft $ 1.00 72 $ 72 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and 

 /

Consumables

Initial CO2 41629 lb 50 lb $ 35.00 833 29,155 $ Cost based on average fertilizer use for Aquaculture and average fertilizer costs from the 

Initial Nutrients 2 lb 1 lb $ 0.20 2 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 7991266 gal 1 gal $ 7583248.679 12,626 0.00 $ H2A Feedstock Costs

Initial Acetate 5442 kg 1 kg $ 0.60 5442 3,238 $ 

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 8,000 $ 

 p q 

http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 3,000.00 $ Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 38 $ 5,548 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 1,500.00 $ Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 4,299.00 $ Information from Emerson Process Management

Level Indicators 1 $ 714.00 38 27,132.00 $ Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 38 13,110.00 $ Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 19 144,400.00 $ Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/anal

Air Temperature Meter 1 $ 599.00 1 599.00 $ Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 38 1,444.00 $ Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 38 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 38 16,530.00 $ Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 38 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Instrument Wiring 213180 ft 1 ft $ 0.02 213180 4,135.69 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 264360 ft 1 ft $ 0.02 264360 5,128.58 $ Information from Waytek Inc. at waytekwire.com

Conduit 8820 ft 1 ft $ 0.58 8820 5,113.84 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 4,274,085

28.2 Photobiological-Fermentor Integration

In the current photobiological pathway configurations, algae and bacteria filtered out of the 

photobiological systems by the rotary drum filters are discarded as waste. However, as 

analyzed in Part II, this waste biomass can be utilized in a fermentation reactor to generate 

additional hydrogen. This integration combines a photobiological pathway in series with a 

fermentation pathway, thus little capital cost efficiencies are expected. Due to the small 

additional hydrogen generated from the fermentation system, it is possible to route the 

fermentation hydrogen produced into the gas capture subassembly of the photobiological 

system to improve costs slightly. As mentioned in Part II, the major cost component of the 

stand-alone algae fermentation systems is the labor cost, which is enormous relative to the 

minimal hydrogen produced. However, the minimal size of the fermentation unit will require 

no additional labor from that already needed for the photobiological plant, leading to massive 

cost savings relative to the stand-alone fermentation system. The process diagram for the 

photobiological-fermentor integration is shown in Figure 28-5. 
174 

Figure 28-5. Photobio-Fermentor Process Diagram

Photobiological-Fermentor Integration LEGEND

Reactor Bed Primary System

Gas Capture Subassembly

Recycle Subassembly

Control System Subassembly

Organism Feed Subassembly

Algal

Fermentation

Hydrogen

and CO2

Acetate

Photo Bio Reactor Bed

Mixers

Power

Source

Control Room PC

Instrumentation

L

T

H2

O2

P

pH

H2

T

Wiring

8 signal

12 power

Raceway

Wiring Panel

Instrumentation

(x 20 Raceways)

Paddlewheel Motors

(x 20 Raceways)

Compressor

Hydrogen

Pipeline Gas

Separation

Gas Capture Subassembly

Hydrogen

Flow Meter

Wiring Panel

Transfer Out

Pump

Wet Spent

Organisms

Makeup

Water

Algal

Separation

Recycle

Subassembly

CO2

Cylinder

Nutrient / Fertilizer

Tank

Metering

Pump

CO2

Saturation

Vessel

Organism Feed

Subassembly

Mixing Pump

Transfer In

Pump

Nutrient

Flow Valve

Water Flow

Valve

Our analysis combines the waste organisms of the B-1, B-2 and B-5 systems with fermentation. 

Since the capital costs and the reactor specifications for B-1 are the same as B-2, we have 

assumed the results for both systems to be the same. Neither the B-3 nor the B-4 systems 

produce any waste algae, since all the algae mass is needed for respiration to keep the systems 

anaerobic. Additionally, the algae in the B-4 system are mounted on an immobilized mat, and 

its waste organisms would to require additional processing which would not result in an 

economically beneficial integration. The B-5 system is similar conceptually to the B-1 and B-2 

system with the exception that it is able to use the waste acetate stream from fermentation to 

feed the photosynthetic organisms. Because of this additional requirement there is an acetate 

retrieval pump required, making the B-5 Photobio-Fermentor Integration a separate integration 

pathway. The Bill of Materials for the B-5 with Fermentation is shown as an example in Figure 

28-6. The Bill of Materials for B-1and B-2 with fermentation would be similar. 
175 

Figure 28-6. BOM for B-5/C-5 Integrated System Description Size Req'd Units Unit Size Units Unit cost Qty Req'd Total Cost Material / Part Description

Photo Bio Reactor Bed Subassembly

Transparent Film 235116 m2 1 m2 $ 0.54 235,116 $ 126,551 Polyethylene price quote at $0.54/m2 ($0.05/ft2)

Pond Lining 253795 m2 1 m2 $ 0.47 253,795 $ 119,912 Butyl rubber, PVC, and LDPE (low density polyethylene)based on quote of $175.56 for 4,000ft2.

Pond Edging 37752 m 1 m $ 7.00 37,752 $ 264,266 Unit Cost is engineering estimate at $7/m

Installation of Ponds 1 raceway $ 26,083.00 54 $ 1,408,482 Based on California labor rates.

Paddlewheel Mixers 1 each $ 5,000.00 108 $ 540,000 Est. paddlewheel cost. 4 per raceway. 18 raceways.

Inlet Water Valve 0.5 in 1 each $ 43.46 54 $ 2,347 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Water Valve 0.5 in 1 each $ 43.46 54 $ 2,347 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Outlet Gas Valve 1.0 in 1 each $ 67.23 54 $ 3,630 Price from ValveStore.com (http://www.valvestore.com/prodinfo.asp?number=551032)

Flanges 0.5 in 1 each $ 8.00 162 $ 1,296 Price from http://www.ancorp.com/line.aspx?id=819

Gas Capture Subassembly

Compressor 45 kgmol/hr 45 kgmol/hr $ 9,233.00 1 414,241 $ Using H2A Unit cost

PSA $ 26,895 1 30,542 $ Using H2A scaling

Raceway Collection Pipe 108 ft 1 ft $ 1.00 108 $ 108 Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

altered) and 2 to 1 molar ratio of hydrogen to oxygen. Pricing from PVC Plastic Corp. 

Main Collection Pipe 3120 ft 1 ft $ 6.18 3120 19,282 $ Sized for Peak Hydrogen production. Assuming Plastic Gas piping, 100 ft/s velocity, uniform pipe diameter (can be 

altered) and 2 to 1 molar ratio of hydrogen to oxygen. Pricing from PVC Plastic Corp. 

Gas Capture Pipe 50 ft 1 ft $ 6.18 50 $ 309 y g p g p p g, / y, p p ( 

altered) and 2 to 1 molar ratio of hydrogen to oxygen. Pricing from PVC Plastic Corp. 

Organism Feed Subassembly

Transfer In Pump 150 gpm 150 gpm $ 10,252 3 30,755 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics

Nutrient Metering Pumps 1468 Gal/Hr. $ 2,594 1 2,594 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics

Mixing Pump 150 gpm 150 gpm $ 10,252 3 30,755 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics

Nutrient Tank 21 gal 10 gal $ 30.00 3 $ 90 10 gallon Cylindrical Process Tank, Part #0275-085 (http://www.watertanks.com/products/0275-085.asp)

Raceway Slurry Collection Pipe 108 ft 1 ft $ 1.00 108 $ 108 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Main Slurry Collection Pipe 3120 ft 1 ft $ 4.31 3120 13,447 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Recycle-Feed Transfer Pipe 180 ft 1 ft $ 6.18 180 1,112 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Consumables

Initial Nutrients 1 lb 1 lb $ 1.112714485 0.20 $ 0 NFC Company, Chicago, IL (800) 734 4515

Initial Water 5720312 gal 1 gal $ 5720311.865 9,524 0.00 $ H2A Feedstock Costs

Initial Acetate 7793 kg 1 kg $ 7792.757824 4,637 0.60 $ Chemical Journal of Korea. Assumes consumption to be four times more to account for full growth conditions.

Recycle Subassembly

Rotary Drum filter 25,741 gph 3750 gph $ 87,000.00 7 609,000 $ Information from Dana Kent at Advanced Aquaculture Inc. on Hydrotech Drum Filter

Transfer Out Pump 150 gpm 150 gpm $ 10,252 3 30,755 $ Est. Based on Perry's Chemical Handbook, Plant Design and Economics

Main Slurry Feed Pipe 3120 ft 1 ft $ 4.31 3120 13,447 $ Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Raceway Slurry Feed Pipe 108 ft 1 ft $ 1.00 108 $ 108 Pricing based off of Estimates from PVC Plastic Corp. Assumes 10 ft/second flow rate and Density of 1g/ml

Fermentation Subassembly

Tank 6173 gal 1 gal $ 2.00 24,693 $ 49,386 This is a scaled cost based on the capital costs from NREL Report TP-510-32438

Slurry Piping - 0.5" 500 ft 1 ft $ 0.52 500 $ 260 Sized using Continuity Equation. Assumed a velocity of 1 ft/sec. 

Slurry Piping - 1" 5077 ft 1 ft $ 1.00 5,077 $ 5,077 Sized using Continuity Equation. Assumed a velocity of 1 ft/sec. 

Pump 6173 gal/day 4 gal/min $ 198.50 1 $ 199 http://www.grainger.com/Grainger/items/1P795

Inlet Valves 1 in 1 each $ 67.23 4 $ 269 http://www.valvestore.com/products.asp?dept=1485

Outlet Valves 0.5 in 1 each $ 46.85 4 $ 187 http://www.valvestore.com/products.asp?dept=1485

Control System

Control Room 160 ft2 1 ft2 $ 50.00 160 8,000 $ comes from price quote of $50/ft2 from http://www.buyerzone.com/industrial/modular_buildings/prefab_guide.html

Control Room Wiring Panel 1 $ 3,000.00 1 3,000.00 $ Consultation with Innomation Systems Inc. 

Raceway wiring Panel 1 $ 146.00 54 $ 7,884 Information from Tessco Technologies Inc. 

Computer and Monitor 1 $ 1,500.00 1 1,500.00 $ Information from Emerson Process Management

Labview Software 1 $ 4,299.00 1 4,299.00 $ Information from Emerson Process Management

Level Indicators 1 $ 714.00 54 38,556.00 $ Omega Engineering Inc. LVR51 Liquid Level Float Transmitter + DPi32 Meter

Pressure Sensors 1 $ 345.00 54 18,630.00 $ Omega Engineering Inc. PX209 Pressure Trnasducer + DPi32 Meter

Hydrogen Area Sensors 1 $ 7,600.00 27 205,200.00 $ Honeywell 7866 Gas Analyzer (http://www.lesman.com/unleashd/catalog/analytical/analyt_hwhydrogengas.htm)

Air Temperature Meter 1 $ 599.00 1 599.00 $ Omega DP81T Meter

Air Temperature Indicator 1 $ 38.00 54 2,052.00 $ Omega TH-10-44000 Thermistor Probe

Water Temperature Indicator 1 $ - 54 $ - Comes with the PH sensor

pH level Indicator 1 $ 435.00 54 23,490.00 $ Omega Engineering Inc. PHE - 6510 PH electrode + DP24 -PH Meter

Oxygen Area Sensors 1 $ - 54 $ - Comes with H2 Sensor

Nutrient Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Water Flow Valve 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Hydrogen Flow Meter 1 $ 5,500.00 1 5,500.00 $ Information from Emerson Process Management

Instrument Wiring 406380 ft 1 ft $ 0.02 406380 7,883.77 $ Information from Waytek Inc. at waytekwire.com

Power Wiring 505080 ft 1 ft $ 0.02 505080 9,798.55 $ Information from Waytek Inc. at waytekwire.com

Conduit 12340 ft 1 ft $ 0.58 12340 7,154.73 $ Information from Waytek Inc. at waytekwire.com

System Initial Cost $ 4,089,570

28.3 Lignocellulosic Fermentation/MEC Integration

The output from a lignocellulosic fermentor would be primarily a mix of acetic acid and ethanol 

with lesser amounts of other organic acids and glucose. NREL fermentor research on the corn 

stover H2 fermentor resulted in the following byproduct: 

acetic acid 26.0 mM

succinic acid 5.6 mM

lactic acid 1.8 mM

formic acid 0.6 mM

 ethanol 14.0 mM
176 

Subsequently, research at Penn State operated the MEC on this product mix. 99 

The fermentation integration evaluated was that of lignocellose fermentation with MEC. This 

integration occurs in series providing little overlap in equipment. The process diagram is shown 

in Figure 28-7. 

Figure 28-7. Process Diagram for Fermentation-MEC Integration 

Microbial

Electrolysis Cell

(MEC)

Gas

Compressor

90%

Hydrogen

Water

Purification

Water

Seed

Production Storage

Hydrolysis Corn Stover

Processing

Saccharification

& Fermentation

Corn Stover

Feedstock

Acetate

Waste

Water

Water

Gas

Compressor

Integrated Lignocellosic Fermentation with MEC

Hydrogen

and CO2

PSA

PSA

Tail Gas

Tail Gas

The MEC component of this integrated system as defined in Part III has been sized to consume 

100% of the lignocellose fermentation waste acetate and other organics. We have assumed no 

extra labor will be needed for the integrated system. Because the hydrogen volumes are large 

and the output gas compositions are different in each case system, separate gas capture systems 

are used for the fermentor and MEC. The Bill of Materials for the Integrated FermentationMEC is shown in Figure 28-8. 

 

99 Logan, Bruce. “Electrochemically Assisted Microbial Fermentation of Acetate.”ZFH-8-77623-01. NREL. 

Colorado. Janurary 2009. 
177 

Figure 28-8. Bill of Materials for Integrated Fermentation-MEC Description Install Factor Size Req'd Units Unit cost Qty Req'd Total Cost

Corn Stover Prep Subassembly

Bale Transport Conveyor 1.62 $ 400,000 2 $ 800,000

Bale Unwrapping Conveyor 1.19 $ 150,000 2 $ 300,000

Belt Press Discharge Conveyor 1.89 $ 50,000 1 $ 50,000

Shredder Feed Conveyor 1.38 $ 60,000 4 $ 240,000

Truck Scales 2.47 $ 34,000 2 $ 68,000

Truck Unloading Forklift 1 $ 18,000 4 $ 72,000

Bale Moving Forklift 1 $ 18,000 4 $ 72,000

Corn Stover Wash Table 2.39 $ 104,000 2 $ 208,000

Shredder 1.38 $ 302,000 4 $ 1,208,000

Concrete Feedstock-Storage Slab 2.2 $ 450,655 1 $ 450,655

Polymer Feed System 2.28 $ 30,000 1 $ 30,000

Wash Table Pump 3.87 $ 20,000 2 $ 40,000

Wash Water Pump 5.19 $ 15,000 2 $ 30,000

Clarifier Underflow Pump 13.41 $ 6,000 1 $ 6,000

Clarified Water Pump 7.07 $ 15,000 1 $ 15,000

Belt Press Sump Pump 2.92 $ 19,000 1 $ 19,000

Clarifier Thickener 1.51 $ 135,000 1 $ 135,000

Belt Press 1.25 $ 100,000 1 $ 100,000

Magnetic Separator 1.3 $ 10,335 1 $ 10,335

Wash Water Tank 2.8 $ 50,000 1 $ 50,000

Clarifier Thickener Tank 3.04 $ 135,000 1 $ 135,000

Pretreatment/Hydrolysis Subassembly

Hydrozylate Mix Tank Agitator 1.2 $ 28,421 1 $ 28,421

Hydrozylate Mixing Tank 1.2 $ 32,237 1 $ 32,237

Hydrozylate Cooler 2.1 $ 23,266 1 $ 23,266

In-line Sulfuric Acid Mixer 1 $ 2,542 1 $ 2,542

Sulfuric Acid Pump 2.8 $ 8,289 1 $ 8,289

Sulfuric Acid Tank 1.4 $ 9,441 1 $ 9,441

Lime Solids Feeder 1.3 $ 3,900 1 $ 3,900

Lime Unloading Blower 1.4 $ 99,594 1 $ 99,594

Lime Dust Vent Baghouse 1.5 $ 140,707 1 $ 140,707

Lime Storage Bin 1.3 $ 136,370 1 $ 136,370

Hydrolysis Tank Pump 2.8 $ 69,516 1 $ 69,516

Methane Boiler 1.3 $ 200,000 1 $ 200,000

Fermentation Subassembly

Fermentation Tank 1.2 $ 493,391 5 $ 2,466,955

Fermentation Agitator 1.2 $ 19,676 10 $ 196,760

Fermentation Cooler 2.1 $ 6,780 5 $ 33,900

Fermentation Recirculation Pump 2.8 $ 13,666 5 $ 68,330

Lignin Wet Cake Screw 1.4 $ 16,837 5 $ 84,185

Acetate Product Storage Tank 1.4 $ 160,829 1 $ 160,829

Acetate Transfer Pump 1.3 $ 53,737 1 $ 53,737

Pneumapress filter 1.4 $ 1,285,736 4 $ 5,142,942

HX - 1 1.3 $ 22,346 1 $ 22,346

HX - 2 1.3 $ 194,618 1 $ 194,618

FTC 1.3 $ 6,781 5 $ 33,905

Seed Production Subassembly

Seed Hold Tank 1.2 $ 160,829 1 $ 160,829

Seed Hold Tank Agitator 1.2 $ 12,492 1 $ 12,492

Seed Hold Transfer Pump 1.4 $ 22,050 1 $ 22,050

1st Seed Fermentor 2.8 $ 14,700 2 $ 29,400

2nd Seed Fermentor 2.8 $ 32,600 2 $ 65,200

3rd Seed Fermentor 2.8 $ 81,100 2 $ 162,200

4th Seed Fermentor 1.2 $ 39,160 2 $ 78,320

5th Seed Fermentor 1.2 $ 146,549 2 $ 293,097

4th Seed Agitator 1.2 $ 11,645 2 $ 23,289

5th Seed Agitator 1.2 $ 10,291 2 $ 20,582

4th Seed Fermentor Coil 1.2 $ 4,658 1 $ 4,658

5th Seed Fermentor Coil 1.2 $ 28,238 1 $ 28,238

Seed Transfer Pump 1.4 $ 53,737 2 $ 107,474

MEC Subassembly

MEC Tanks 1.3 $ 0.51 1.13E+08 $ 57,488,911

Brush Cathode 1.2 $ 14.60 27436736 $ 400,686,098

Brush Anode 1.2 $ 7.92 7868218 $ 62,316,288

Power Supply 1.3 $ 350,000.00 53 $ 18,550,000

Process Water Circulating Pump 2.8 $ 8,818 1 $ 8,818

MIC - 1 1.3 $ 88,542 1 $ 88,542

MIC - 2 1.3 $ 554,634 1 $ 554,634

Cooling Water Pump/MEC Pumps 2.8 242,216 12 $ 2,906,592

Storage Subassembly

CSL Storage tank Agitator 1.2 2,143 1 $ 2,143

CSL Pump 2.8 6,188 1 $ 6,188

CSL Storage Tank 1.4 61,949 1 $ 61,949

CSL/DAP Day Tank 1.4 28,470 1 $ 28,470

CSL/DAP Pump 2.8 6,188 1 $ 6,188

CSL/DAP Day Tank Agitator 1.2 12,348 1 $ 12,348

DAP Solids Feeder 1.3 3,900 1 $ 3,900

DAP Unloading Blower 1.4 49,014 1 $ 49,014

DAP Vent Baghouse 1.5 9,595 1 $ 9,595

DAP Storage Bin 1.3 34,255 1 $ 34,255

Sulfuric Acid Pump 2.8 13,814 1 $ 13,814

Sulfuric Acid Storage Tank 1.2 60,470 1 $ 60,470

Propane Storage Tank 1.4 30,825 1 $ 30,825

Make-up Water Pump 2.8 8,735 1 $ 8,735

Process Water Circulating Pump 2.8 8,818 2 $ 17,635

Process Water Tank 1.4 168,502 2 $ 337,004

Gas Capture Subassembly

Compressor Fermentor 1.3 1704 kgmol/hr $ 9,233 1,704 $ 15,729,564

FC-1 1.3 $ 29,623 1 $ 29,623

FIC -1 1.3 $ 57,651 1 $ 57,651

FIC -2 1.3 $ 311,057 1 $ 311,057

Compressor MEC 1.3 1929 kgmol/hr $ 9,233 1929 $ 17,812,305

PSA Fermentor 1.3 1 $ 1,527,494

PSA MEC 1.3 1 $ 302,154 Hydrogen Flow Meter 1.3 $ 5,500 1 $ 5,500.00

System Initial Cost $ 593,317,371
178 

28.3.1 Flow Chart

The combined system mass flows and flow properties are diagramed in Figure 28-9. Details of 

these components and the reactions are included in the Part II discussions. A detailed Mass and 

energy balance is included in Part IV, Appendix A. 

Figure 28-9. Fermentor/MEC Mass Balance Flow Chart

Fermentation and MEC Integrated System

Efficiencies- Hydrolysis: 90%, Fermentation: 90%, MEC: 90%

HX-2 HX-1 Feed

 Heater 416,667kg/hr +85C pump 416,667kg/hr +23C 416,667kg/hr Stock

11,214 kW 128C 4.7 atm 43C 1 atm 20C Processors

 150C -85C -25C 20C

 4.7 atm 4.7 atm expander 1 atm 30C 20% solid

416,667kg/hr 388,860kg/hr Lignin 80% water

Hydrolyzer Saccharification Separator

416,667kg/hr Liquid mash 72,199kg/hr

416,667kg/hr 55C HX 388,860kg/hr

1 atm Fermentor acetate 27,403kg/hr

150C in Fermentation Temperature glucose 4,567kg/hr

140C out 55C Control water 284,691kg/hr

4.7 atm Lime 1 atm

416,667kg/hr Gas: MEC unreacted liquids waste

H2 1,937kg/hr Reactors

CO2 21,139kg/hr CO2

 H20 4,730kg/hr Water 15,252,803kg/hr

Condenser FC-1 Supply 30C

14,936,142kg/hr 1 atm H2 3,863kg/hr

1 atm 27,806kg/hr water CO2 2,109kg/hr

14.7 psi 2,044kg/hr H2O 1,531kg/hr

MEC

Compressor

Fermentor Compressor Intercooler stage 1

stage 1 FIC-1 60 psi Intercooler

60 psi 25,762kg/hr 7,503kg/hr MIC-1 liquid water

25,762kg/hr water stage 2 7,503kg/hr

Fermentor Compressor 45C 2,054kg/hr 302 psi 45C, 60 psi High purity H2

stage 2 60 psi 6,834kg/hr PSA

302 psi 23,708kg/hr Intercooler 6,135kg/hr

Intercooler H2 PSA High purity H2 MIC-2

 FIC-2 CO2 80% 6,834kg/hr CO2+H2O 2,272kg/hr

45C waste H2 193kg/hr

23,708kg/hr 23,196kg/hr 302 psi water

45C 40C CO2+H2O 21,259kg/hr

302 psi H2 387kg/hr Waste H2

water

512kg/hr

3,670kg/hr

12kg/hr

40,064kg/hr

92,721 kg/day

698kg/hr

88,085 kg/day

Storage

15,205,236kg/hr

670kg/hr

water

Corn

Stover

Feed Stock

Water

20,440kg/hr

VH2O

Excess H2

116kg/hr

 H2O

375kg/hr

with:

 CO2

Water

Separation

and Recovery

46,477 kg/day

37,181 kg/day

1,549kg/hr

29.

The hydrogen production costs for the previously described system combinations are discussed 

below. By comparing the hydrogen production costs from the integrated systems to the 

individual systems we can determine if there is any benefit to systems integration. A cost 

comparison for the integrated systems is shown in 

Results and Discussion for Integrated Systems

Figure 29-1. 
179 

Figure 29-1. Integrated Systems Results

Capital Costs H2 Production 

(kg/day) H2 Cost ($/kg)

Stacked Photobio Beds

Stand alone B-3 $39,485,240 10,000 $4.17

Stand alone B-5 $34,196,970 10,000 $10.36

Integrated $41,485,500 10,600 $5.25

Photobio-Fermentor Integration

Stand Alone B-1/B-2 $21,880,470 10,000 $2.99

Stand Alone Algae Fermentation (C1/C2) $37,142 7 $172.73

Integrated $22,272,680 10,007 $3.21

Photobio-Fermentor Integration

Stand Alone B-5 $34,196,970 10,000 $10.36

Stand Alone Algae Fermentation (C5) $86,211 13 $66.17

Integrated $34,591,120 10,013 $11.04

Lignocellulosic Fermentor and MEC

Lignocellulosic Alone $44,944,078 37,181 $4.33

MEC Alone $558,689,165 88,055 $12.43

Integrated $593,317,371 125,266 $6.61

Systems

The cost change between the stand-alone systems and the integrated systems is shown in 

Figure 29-2. 

Figure 29-2. Comparison of Integrated System Costs

$172.73 $66.17

$0.00

$2.00

$4.00

$6.00

$8.00

$10.00

$12.00

$14.00

$16.00

$18.00

$20.00

Stand alone B-3 Stand alone B-5 Integrated Stand Alone B-1/B-2 Stand Alone Algae 

Fermentation (C1/C2)

Integrated Stand Alone B-5 Stand Alone Algae 

Fermentation (C5)

Integrated Lignocellulosic Alone MEC Alone Integrated

H2 Cost/kg

System

Comparison of Integrated Systems
180 

30.

Integrating hydrogen producing systems can yield greater hydrogen production per area of land, 

however, it is not necessarily the most cost effective method of producing hydrogen for the 

following reasons: 

Integrated System Conclusions and Recommendations

• The increase in capital costs for the stacked photobioreactor beds and the 

photobioreactor plus fermentor, while less than the systems alone, are still substantial 

enough to outweigh the increase in hydrogen output. 

• For the Fermentor/MEC combination, the high cost of H2 from the MEC makes the 

combination uneconomic. This is partly due to the current immature and unoptimized 

status of the MEC, and the results could be markedly improved with additional MEC 

developments and cost optimization. 
181 

31.

Processes Mass flow mole flow T T ave ΦT P Cp h dq/dt

% kg/day kg/hr kgM/hr °C °C °C psi kJ/kg °C kJ/kg kJ/hr

Feedstock Corn stover 20.0% 2,000,000 83,333 20 14.7

Inputs cellulose 36.7% 734,000 30,583

hemicell 27.4% 548,000 22,833

Lignin,etc 35.9% 718,000 29,917

total 100.0% 2,000,000 83,333

water 80.0% 8,000,000 333,333 20 14.7

total slurry 100.0% 10,000,000 416,667 20 14.7

HX-1 from feedstock 10,000,000 416,667 20.0 14.7 4.180

to HX-2 10,000,000 416,667 43.3 31.7 23.3 40,655,319

from fermentor 9,332,641 388,860.06 55.0 14.7 4.182

to LigSep 9,332,641 388,860 30.0 42.5 -25.0 -40,655,319

HX-2 from HX-1 primary 10,000,000 416,667 43.3 68.8 4.206

to hydrolyzer heater 10,000,000 416,667 128.5 85.9 85.2 149,281,250

from hydrolyzer 10,000,000 416,667 140.0 85.0 68.8 4.215

to sacchrification 10,000,000 416,667 55.0 97.5 -149,281,250

Hydolyzer heating of slurry from HX-2 10,000,000 416,667 128.5

Heater 10,000,000 416,667 150.0 139.3 21.5 68.8 4.290 38,387,188

heating of CO2 from PSA 490,873 20,453 45.0

CO2/H2mass 54.58 490,873 20,453 150.0 97.5 105.0 14.7 0.913 1,960,729

heating of H2O from PSA 2,783 116 45.0

H2O/H2mass 0.31 2,783 116 150.0 97.5 105.0 14.7 1.970 23,985

Total heat input required (neglecting H2 ΦT) 40,371,902

Waste H2 burner heating effic =

H2 LHV kJ/kg = 119,705 90% 8,994 375 14.7 40,371,902

 10,000,000 416,667

Total heating from HX-1 + HX-2 + Heater 230,308,471

Hydrolyzer In from Heater 10,000,000 416,667 150.0 68.8

inputs cellulose 36.7% 734,000 30,583 4.68

hemicell 27.4% 548,000 22,833 atm

Lignin,etc 35.9% 718,000 29,917

water 80.0% 8,000,000 333,333

total 10,000,000 416,667

efficiency= output complex sugars 1,153,800 48,075 140.0

90% xylan 128,200 5,342

Lignin,etc 718,000 29,917

organics 2,000,000 83,333

water 8,000,000 333,333

total 10,000,000 416,667

Endothermic heat absorption ΦT 10.0

to HX-2 10,000,000 416,667 140.0 68.8

Saccharification from HX-2 10,000,000 416,667 55.0

in fermentor to fermentation 10,000,000 416,667 55.0

Fermentor in from total 10,000,000 416,667 55.0

efficiency = sacchrification simple sugars 1,153,800 48,075 267

(glucose)

90% Lignin,xylan,etc 846,200 35,258

water 8,000,000 333,333

reactants 10,000,000 416,667

H2O reacted in fermentation 207,676 8,653

gas out H2 46,476.6 1,937 961 55

CO2 507,339.5 21,139 480 55

V H2O 113,542 4,731 263 55 14.7

total gas 667,359 27,807 1,704

slurry out to HX-1 slurry total 9,332,641 388,860 55

Lignin,xylan,etc 846,200 35,258

Acetate 692,280 28,845 

Glucose 10% 115,380 4,808 

L-H20 7,678,781 319,949

total 9,332,641 388,860

total mass out 10,000,000 416,667

Condenser gases in H2 46,477 1,937 55 14.7 14.496

FC-1 from fermentor CO2 507,340 21,139 55 14.7 0.953

V H2O 113,542 4,731 55 14.7 1.912

total 667,359 27,807

gases out to H2 46,477 1,937 960.6 45 50 -10 14.7 14.496 -280,719

Compressor-F-1 CO2 507,340 21,139 480.3 45 50 -10 14.7 0.953 -201,350

V H2O 64,477 2,686.5 149.1 45 50 -10 14.7 1.912 -51,367

total gas 618,293 25,762 -533,436

liquids out L H2O 49,066 2,044 45 -10 14.7 4.200 -62,477

Part IV Appendix A: Integrated Fermentor/MEC Mass/Heat Balance
182 

Processes Mass flow mole flow T T ave <T P Cp h dq/dt

% kg/day kg/hr kgM/hr °C °C °C psi kJ/kg °C kJ/kg kJ/hr

Compressor in H2 46,477 1,937 961 45 14.7

stage 1 CO2 507,340 21,139 480 45 14.7

F-1 V H2O 64,477 2,687 149 45 14.7

total 618,293 25,762 1,590 45 14.7

out H2 46,477 1,937 961 190.8 60

CO2 507,340 21,139 480 190.8 60

V H2O 64,477 2,687 149 190.8 60

total 618,293 25,762 1,590

Intercooler Gas in H2 46,477 1,937 961 190.8 60 14.496

FIC-1 CO2 507,340 21,139 480 190.8 60 0.953

 Primary V H2O 64,477 2,687 149 190.8 60 1.912

total 618,293 25,762

Gas out H2 46,477 1,937 960.6 45 117.9 -146 58 14.496 -4,092,881

CO2 507,340 21,139 480.3 45 117.9 -146 58 0.953 -2,935,688

V H2O 15,169 632.1 35.1 45 117.9 -146 58 1.912 -176,199

total gas 568,986 23,708 -7,204,769

Liquid out L H2O 49,307 2,054 45 -146 58 4.200 -915,401

V-L Phase change heat -2222 -4,565,040

total out 1,187,279 49,470 -12,685,209

 Secondary water in 3,642,123 151,755 20 4.180

water out 3,642,123 151,755 40 30 20 4.180 12,685,209

Compressor in H2 46,477 1,937 45 58

stage 2 CO2 507,340 21,139 45 58

F-2 V H2O 15,169 632 45 58

total 568,986 23,708 1,476

out H2 46,477 1,937 960.6 218 302

CO2 507,340 21,139 480.3 218 302

V H2O 15,169 632 35.1 218 302

total 568,986 23,708 1,476.0

FIC-2 Gas in H2 46,477 1,937 218 302

 Primary CO2 507,340 21,139 218 302

V H2O 15,169 632 218 302

total 568,986 23,708

Gas out H2 46,477 1,937 960.6 45 131.45 -173 300 14.496 -4,853,629

CO2 507,340 21,139 480.3 45 131.45 -173 300 0.953 -3,481,348

V H2O 2,876 120 6.7 45 131.45 -173 300 1.912 -39,619

total gas 556,692 23,196 -8,374,596

Liquid out L H2O 12,293 512 45 -173 300 4.200 -270,646

V-L Phase change heat -2222 -1,138,145

total out 568,986 23,708 -9,783,386

 Secondary water in 2,808,964 117,040 20 4.180

water out 2,808,964 117,040 40 30 20 4.180 9,783,386

PSA in total 556,692 23,196 45 300

 fermentor H2 46,477 1,937 960.6 45 300

CO2 507,340 21,139 480.4 45 300

V H2O 2,876 120 6.7 45 300

out H2 outlet 80% 37,181 1,549 45 300

H2 loss 20% 9,295 387 192.1 45 300

CO2 507,340 21,139 480.4 45 300

CO2/H2 mass ratio 54.6

V H2O 2,876 120 6.7 45 300

VH2O/H2 mass ratio 0.31

waste gas out 519,511 21,646 679.2

Burner use of waste H2= 97%

LigninSeparator in from HX-1 9,332,641 388,860

(from fermentor) Lignin, etc 846,200 35,258 30 14.7

Acetate 692,280 28,845

Glucose 115,380 4,808 

L-H20 7,678,781 319,949

total 9,332,641 388,860

out - mash Lignin, etc 100% 846,200 35,258 30 14.7

Acetate 5% 34,614 1,442

Glucose 5% 5,769 240

LH2O/lignin 100%

LH2O 846,200 35,258

total 1,732,783 72,199

out - liquid Acetate 95% 657,666 27,403 30 14.7

to MEC Glucose 95% 109,611 4,567

H20-L 6,832,581 284,691

total 7,599,858 316,661
183 

Processes Mass flow mole flow T T ave θT P Cp h dq/dt

% kg/day kg/hr kgM/hr °C °C °C psi kJ/kg °C kJ/kg kJ/hr

MEC water addition 339,082,263 14,128,428

MEC from LigSep Acetate 657,666 27,403 456.3 30

0.9V system Glucose 109,611 4,567 25.4

total substrate 767,277 31,970 481.7

H20-L 6,832,581 284,691

Water addition: total input 7,599,858 316,661

 Acetate H2O + substrate= 346,682,121 14,445,088

dilution (g/L) = Total H20 need= 345,914,844 14,413,119

90% 2.213 Water Addition 339,082,263 14,128,428

reaction total water 345,914,844 14,413,119

completion total input 346,682,121 14,445,088

reactant vol.=

reacted acetate+glucose 90% 690,549 28,773

reacted water 414,314 17,263 958.3

output-gas H2 92,721 3,863 1,916.5 30 14.7

calculation CO2 total 1,012,142 42,173 958.3 30 14.7

completion CO2 in H2 5% 50,607 2,109 47.9 0.05001

(process chart) H2O-V 36,752 1,531.3 85.0 30 14.7

total cathode gas 1,141,615 47,567 2,049 30

output-liquid waste Acetate 10% 65,767 2,740 30

Glucose 10% 10,961 457 30

H20-L 345,463,778 14,394,324

total liquids 345,540,506 14,397,521

total output 346,682,121 14,445,088

CO2 Vented from MEC 95% 961,535 40,064

Compressor in H2 92,721 3,863 1,916 30 14.7

MC-1 CO2 5% 50,607 2,109 48 30 14.7

V H2O 36,752 1,531.3 85.0 30 14.7

total 180,080 7,503 2,049 30 14.7

out H2 92,721 3,863 1,916 204 60

CO2 50,607 2,109 47.9 204 60

V H2O 36,752 1,531 85 204 60

total 180,080 7,503 2,049

Intercooler Gas in H2 92,721 3,863 204 60

MIC-1 CO2 50,607 2,109 204 60

 Primary V H2O 36,752 1,531 204 60

total 180,080 7,503

Gas out H2 92,721 3,863 1,916 45 124.3 -159 58 14.496 -8,882,141

CO2 50,607 2,109 47.9 45 124.3 -159 58 0.953 -318,543

V H2O 20,679 861.6 47.8 45 124.3 -159 58 1.912 -261,287

total 164,007 6,834 -9,461,971

Liquid out L H2O 16,072 670 45 -159 58 4.200 -324,584

V-L Phase change heat -2222 -1,488,040

total out 180,080 7,503 -11,274,595

 Secondary water in 3,237,113 134,880 20 4.180

water out 3,237,113 134,880 40 30 20 4.180 11,274,595

Compressor in H2 92,721 3,863 45 58

MC-2 CO2 50,607 2,109 45 58

V H2O 20,679 862 45 58

total 164,007 6,834 2,012 45 58

out H2 92,721 3,863 1,916 267 302

CO2 50,607 2,109 48 267 302

V H2O 20,679 862 48 267 302

total 164,007 6,834 2,012

Intercooler Gas in H2 92,721 3,863 267 302

MIC-2 CO2 50,607 2,109 267 302

 Primary V H2O 20,679 862 267 302

total 164,007 6,834

Gas out H2 92,721 3,863 1,916 45 156 -222 300 14.496 -12,432,757

CO2 50,607 2,109 48 45 156 -222 300 0.953 -445,880

V H2O 3,921 163.4 9.1 45 156 -222 300 1.912 -69,347

total gas 147,249 6,135 -12,947,984

Liquid out L H2O 16,758 698 45 -222 300 4.200 -473,726

V-L Phase change heat -2222 -1,551,548

total out 164,007 6,834 -14,973,258

 Secondary water in 4,299,057 179,127 20 4.180

water out 4,299,057 179,127 40 30 20 4.180 14,973,258

PSA in H2 92,721 3,863 1,914 45 300

 MEC CO2 50,607 2,109 48 45 300

91,524,080 gallons
184 

Part V: Discussion of Project Results
185 

32. Summary of Results and Conclusions

This report consists of technoeconomic boundary analyses conducted for thirteen different 

biological hydrogen production systems. Eight were stand-alone systems:

• Five photosynthesis systems using algae and bacteria 

• A fermentation system for waste algae

• A fermentation system for lignocellulose

• A Microbial Electrolysis Cell (MEC)

Five integrated systems were examined, each combining two of the above systems: 

• A vertically stacked photosynthesis system using algae and PNS bacteria

• Three systems combining photobiological and fermentor systems

• A lignocellulose fermentation system combined with an MEC to process the 

organic byproduct of the fermentor 

32.1 Photosynthesis systems

Each photosynthesis system used a series of large ponds or beds for organism growth and 

subsequent hydrogen production. The reactor beds were covered with a non-porous transparent 

cover to capture the H2 and other gaseous products. Multiple configuration approaches were 

considered including single bed, dual bed, and chemostat configurations. The best single 

configuration for reach system was selected for more detailed study. Cost analysis suggests that 

large shallow reactor beds with thin film LDPE covers are the most economic approach. Such 

reactor configurations were selected for all photobiological H2 production. 

The five photosynthesis systems were: 

• B-1: Chlamydomonas reinhardtii mutant with O2 tolerant hydrogenase 

• B-2: Cyanobacterium mutant with O2-tolerant hydrogenase 

• B-3: Chlamydomonas reinhardtii with mutated sulfate permease

• B-4: Chlamydomonas reinhardtii, immobilized on a porous mat and sulfur 

deprived 

• B-5: Rhodobacter sphoeroides RV Purple Non-Sulfur (PNS) Bacteria

The B-1, B-2 and B-5 systems were used in a chemostat II configuration, which is able to 

continuously generate H2 while the organisms are regenerated using a portion of the solar 

radiation. The B-3 system alternated between H2 generation (3 days) and algae regeneration (4 

days). The B-4 system was immobilized on a polypropylene film, thus minimizing the algae’s 

energy requirements. This system operated cyclically, alternating 3 days production and 1 days 

regeneration. 

For the hydrogen production calculations, the following key assumptions were made affecting 

solar-to-hydrogen (STH) energy efficiency for the upper bound of organism functioning: 

1. PAR (Photosyntheticaly Active Radiation): 44% for B-1 through B-4 and 71% 

for B-5 
186 

2. Antennas: Mutants are developed with highly truncated antennas for maximal 

solar utilization (mutants have been developed with moderately truncated 

antennas but these have not produced significant amounts of hydrogen) 

3. Light saturation: Mutants are developed that do not have ETR (Electron Transfer 

Rate) limitations in sunlight up to maximum intensity, thus allowing 12.2% STH 

energy efficiency. (Current mutants are saturated at low sunlight levels, limiting 

STH efficiency to 2-3%)

4. Chemostat II: A continuously operating Chemostat II is feasible, allowing 

simultaneous growth and H2 production. (This continuous operation has not yet 

been demonstrated in comprehensive experiments)

Near term STH energy conversion efficiencies were also estimated for the photobiological 

systems. The key assumptions listed above for the upper bound cases generally apply with the 

exception of an imposition of light saturation affects (ie. ETR limitations). These near term 

efficiencies are somewhat subjective and are higher than performance currently demonstrated in 

the lab but much lower than upper bound performance. 

Figure 32-1summarizes the solar-to-hydrogen energy conversion efficiency assumed for both 

upper bound performance and near term performance.

Figure 32-1. Photobiological System STH Energy Efficiency100

STH Efficiency B1 B2 B3 B4 B5 

Upper Bound 9.2% 9.2% 5.2 % 2.25% 3.5%

Near Term 2% 2% 1.3% 1.5% 1.5%

32.2 Algae Fermentation Systems

Fermentation systems were designed to carry out dark fermentation of waste organisms as 

would be generated in the B-1 through B-5 systems. The maximum potential output from 

glucose is 4 moles H2 per mole glucose. The fermentation converts the organisms into H2, CO2, 

acetic acid, and other organic compounds. The algae fermentation process has been 

experimentally demonstrated at lab scale by NREL using NREL-developed bacteria consortia. 

These initial tests at NREL partially converted cell starches, and subsequently, using the same 

consortium, additionally converted lipids. The high starch algae (sulfur deprived for 21 hours 

and 32% glucose) yielded 1.7 moles H2 per mole glucose, or H2 mass of 0.618% of algae mass. 

The low starch algae (sulfur deprived for 142 hours and 6% glucose) yielded 2.3 moles H2 per 

mole glucose, or H2 mass of 0.155% of algae mass. With additional experimental 

developments, near term future mass conversion rates are expected to be 0.4wt% from sulfur 

depleted algae.

For this study, the fermentor unit was sized based on the waste algae output from the 

photosynthesis systems. With this relatively small feedstock, the ferment output was low, 

resulting in high H2 production cost. 

 

100 STH Efficiency =Solar to Hydrogen conversion efficiency = ratio of hydrogen net energy produced (lower

heating value) to total solar energy incident on reactor bed.
187 

For the algae fermentor production calculations, the following key assumptions were made:

1. Mass conversion rate of sulfur deprived algae was 0.4% H2 mass/algae mass. 

(Current test achievements include: 0.154% for sulfur deprived 142 hrs and 

0.613% for sulfur deprived 21 hrs) 

Only an upper bound performance estimate was made for the fermentation system since the 

system is at a much more developed level than the photobiological system.

32.3 Lignocellulose Fermentation System

The lignocellulose fermentation system used corn stover as a feedstock to generate hydrogen 

via a dark fermentation process. The process is quite complex and involves extensive feedstock 

pre-processing, hydrolysis, saccharification, and fermentation. Portions of the process design 

and capital costs were based on an NREL corn stover-to-ethanol plant cost analysis. The unique 

150oC hydrolysis process used for this study is one that has been demonstrated by NREL. The 

unique organisms for the saccharification process and for the fermentation-to-H2 process have 

been developed by NREL at lab scale. The saccharification process utilizes organisms to 

convert the hydrolysis products into simple sugars in a short time. The subsequent fermentation 

process uses the same bacteria to produce H2 and CO2 gas plus simple organic liquids 

For the lignocellulose fermentor production calculations, the following assumptions were made: 

1. Acid hydrolysis to break down biomass to hemicellulose and cellulose 

components of complex sugars can be effectively carried out at a temperature of 

150oC and pressure of 4.7 atm. with an efficiency of 90% conversion of the 

components. 

2. Saccharification and fermentation can be effectively carried out sequentially in 

the same reactor vessel at a temperature of 55oC , at atmospheric pressure, and 

using 40 hr batch cycle with an efficiency of 90% conversion. 

The saccharification process utilizes bacteria to convert the hydrolysis products into simple 

sugars in a short time. The subsequent fermentation process uses the same bacteria to produce 

H2 and CO2 gas plus organic liquids, requiring approximately 40 hours. Efficiency for these 

processes in combination was 90% conversion to H2 and simple liquid organic byproducts. 

Costs of 37 tonne/day H2 production were initially calculated assuming zero value of the 

byproducts. These byproducts are essentially: 692 tonnes of organic acid and alcohol liquids 

plus 846 tonnes of lignin and xylan solids. The liquid byproducts comprising 51% acetate and 

21% ethanol have intrinsic value, but would require a subsequent separation process such as 

distillation. Taking credit for the value of the large amount of byproduct would substantially 

reduce net H2 costs. A more rigorous assessment of this effect should be conducted. As an 

alternative to byproduct recovery, the liquid organics can be processed in an MEC to produce 

additional hydrogen. 

In addition to the liquid byproducts, the extracted lignin solids can be used as a low grade fuel 

for industrial processes. 
188 

Like the algae fermentation system, only an upper bound performance estimate was made for

the lignocellulosic fermentation system. 

32.4 Microbial electrolysis Cell (MEC) System

The MEC has been developed at Penn State University to break down simple liquid organics 

such as acetic acid and ethanol mixed with water into H2 and CO2, with the gasses being 

separated by the electrolysis process. 

For the electrolysis process, the anode microbes generate a voltage which is supplemented with 

an external voltage to generate an electrolysis current. A lower external voltage results in 

higher efficiency, but increases reaction times, thus increasing system size and cost. It was 

determined that for lowest cost H2, the system should be operated at a higher than normal 

voltage (0.9V) to minimize capital plant costs. The very low concentration of the input acetic 

acid (~2g acetate/Liter of water as opposed to the fermentor at 200g/L) resulted in a very large 

system, i.e., 96,000,000 gallons of total reactor volume for 88 tonnes H2/day. These volumes 

and the required high area anodes and cathodes were based on specifications from lab tests by

Penn State. The anode used was a carbon brush with the microbe layer and the cathode was a 

stainless steel brush. The details of utilization of these electrodes in a large scale reactor tank 

have not been fully worked out. Since the MEC separates most of the H2 from the CO2, its PSA 

requirements are greatly reduced and there is reduced loss of H2 in the PSA process. 

Key assumptions for the MEC system were:

1. Conversion efficiency of acetate to H2 is 90% of the theoretical 4 moles of H2/mole 

acetate (lab tests have achieved as high as 95%)

2. The lab-scale system design will function equally well for very large scale reactors.

3. Acetate feedstock for the stand-alone MEC was $0.60/kg, (which is a significant 

component of the resultant H2 cost).

Only an upper bound performance estimate was made for the MEC system given its currently 

very low stage of development. 

32.5 Integrated Stacked Photosynthesis System

This system employs two different photosynthesis systems having different, but overlapping, 

PAR wavelength bands. The narrower band system is stacked on top of the wider band system 

so that the lower system can utilize photons not utilized by the upper system. For this case, the 

more efficient (4 photons/H2), but narrower band Chlamydomonas reinhardtii was above the 

less efficient (11 long wavelength photons/H2), but wider band (PAR = 0.71) PNS bacteria. For 

the top bed, the B-3 system was used, since the B-1 and B-2 mixtures of O2 and H2 raised safety

concerns for this application and the B-4 mat and high cell density absorbed too much sunlight. 

The stacked system has the benefit of reduced land area requirement. 
189 

32.6 Integrated Photosynthesis and Fermentor System

This system used the algae or bacteria waste from the photosynthesis beds as feedstock to an 

algal/organism fermentor. The B-1, B-2 and B-5 systems were chosen as the best adapted for 

the integration, since, in chemostat operation, they will be filtering waste algae from a slip 

stream of water to maintain bed concentration. In the B-3 and B-4 system, no waste algae is 

removed due to the need for the system to respire enough glucose to keep the system anaerobic. 

In all cases, the quantity of feedstock was so low that very little H2 was produced as compared 

with the 10 tonne/day photosynthesis beds. For the B-1 and B-2 integrations, output was 

~7kg/day and for the B-5 integration, output was ~18 kg/day. An additional benefit of the B-5 

integration was that the fermentor produced acetic acid byproduct, providing 3% of the B-5 

acetic acid requirement. 

32.7 Integrated Lignocellulosic Fermentor and MEC

The baseline lignocellulosic fermentor and baseline MEC were sized to allow integration of the 

two such that the organic liquid byproduct of the fermentor could be used as the entire feedstock 

for the MEC. This byproduct has been shown in NREL experiments to be 51% acetic acid and 

21% ethanol, with the remainder being other organic acids. Thus, the 2,000 tonnes/day corn 

stover plus water produced 37 tonnes H2/day from the fermentor and 88 tonnes H2/day from the 

MEC. 

32.8 Gas Processing

For each of these systems, the reactor gaseous outputs were ducted to compression and gas 

separation systems. The compressors were two stage units, using intercoolers to reduce gas 

temperature and to condense water vapor, thus minimizing compression power requirement. 

The final separation was carried out with PSA (Pressure Swing Adsorption) units that would 

produce 99.9% pure H2. 

32.9 Hydrogen Production Cost Comparisons

Given the systems and performance assumptions listed above, the feasibility, performance, 

capital cost, and resultant $/kg H2 were evaluated for each system and integrated system. 

System hydrogen production costs are summarized in Figure 32-2. 
190 

Figure 32-2. H2 Production Costs

Near Term 

Performance

Upper Bound 

Performance

Photobiological H2 Production

B-1: Algal O2-tolerant Hydrogenase 10,000 $8.15 $2.99

B-2: Cyanobacterium O2-tolerant Hydrogenase 10,000 $8.15 $2.99

B-3: Algal Sulfate Permease 10,000 $10.48 $4.17

B-4: Immobilized Algal, Sulfer deprived 10,000 $8.44 $6.02

B-5: PNS Bacteria 10,000 $13.95 $10.36

Fermentation of Waste Algae/Photobacteria

C-1 or C-2 (Effluent from B-1 or B-2) 7 -- $172.73

C-5 (Effluent from B-5) 19 -- $66.17

Fermentation of Lignocellulose --

System with no byproduct credit 37,181 -- $4.33

System with byproduct sales ($0.12/kg acetate byproduct) 37,181 $2.09

 --

MEC - Microbial Electrolysis Cell (Acetic Acid Feedstock) 88,055 -- $12.43

Integrated Systems

Integrated Photobiological - Stacked System (B-3 over B-5) 10,600 -- $5.25

Integrated Photobiological/Fermentor

B1-C1 or B2-C2 Integration 10,007 -- $3.21

B5-C5 Integration 10,019 -- $11.04

Integrated Lignocellulosic Fermentor/MEC 125,266 -- $6.61

H Hydrogen Cost, $/kg 2 

Production

kg/day

System

For the pure photobiological systems, the upper bound performance B-1/B-2 system101

 

101 While the B-1 and B-2 pathways utilize different organisms, all system engineering parameters are identical. 

Consequently, predicted hydrogen cost is the same.

achieved 

the lowest hydrogen cost, however, these results are predicated on major improvements in 

organism mutations achieving truncated antenna reductions, elimination of cell light saturation 

due to Electron Transfer Rate (ETR) limits, and satisfactory chemostat II operation. The B-3 

and B-4 system costs are slightly higher and the systems are more complex, but the components 

have been more completely demonstrated. The B-5 system has the highest cost of the 

photobiological systems due the larger amount of photons needed (11 to 15 vs. 4) and the high 
191 

cost of the acetic acid feedstock. This system definition also assumed a satisfactory Chemostat 

II. 

Not surprisingly, the photobiological systems achieving near term STH energy efficiency are 

observed to be substantially higher cost than those achieving upper bound STH efficiency. The 

cost difference is a function of both required reactor area and scaling affects of non-bed 

components, thus the cost difference cannot be determined by simply ratioing STH efficiencies. 

The algae/bacteria fermentation systems have high costs due to the low organism feedstock

input resulting in low H2 output. A large part of the resulting H2 cost is due to the labor costs, 

(94-96% of total cost) since we are analyzing it as a stand-alone system. The cost contribution 

of the labor drops drastically when integrated with the Photobiological system. The current 

results are based on a projected fermentation output of highly sulfur-deprived organisms and do

not exploit fully the organic components of the algae feedstock. It is expected that future 

bacteria and processing developments can facilitate more extensive conversion of the starch, 

lipid and protein content of the algae into H2. Note that lignocellulose was not added to the 

algal fermentor to supplement the feedstock due to different bacterium needed and significant 

transportation costs from lignocellulosic sources to the high solar intensity photobiological sites. 

The lignocellulosic fermentation achieved a moderately low H2 cost, using bacteria and 

processing that has been proven in lab environments, but not in large scale demonstrations. 

There is also high potential for cost reduction from sales of the 51% acetic acid content liquid 

byproduct. If the byproduct had a market value of $0.12/kg (as compared to the market price of 

~ $0.60/kg for acetic acid), the net H2 cost would be reduced to near $2.00/kg. 

For the MEC system, the moderately high H2 cost resulted from the very dilute acetic acid/water 

reactant, and necessitated a very large reactor volume and correspondingly very large anode and 

cathode areas. This high capital cost was coupled with high acetate market price, which could 

potentially be reduced by lowering acetic acid purity, which is not marketed, but is available as 

a fermentor byproduct. The immaturity of the full scale system concepts and components 

indicated that there is extensive potential for future cost reductions. Cost saving could also arise 

from higher concentration of electrolyte and higher pressure operation. 

For the integrated stacked photobiological system, the H2 cost is between the cost of the 

individual stand-alone systems. This indicates that there is no cost benefit to integration as a 

scaled-up version of the lowest cost non-integrated configuration is most preferable. 

For the integrated photobiological algae/fermentor system, the resulting H2 costs are higher than 

the stand-alone photobiological system, also indicated that integration is not preferred. 

For the integrated fermentor/MEC system, the MEC’s H2 production cost is reduced 

significantly due to supply of a “free” feedstock. However, due to high MEC capital costs, the 

cost of the combined system is still significantly higher than the fermentor alone. 
192 

33. Recommendations for Future Work

Key recommendations for future work are summarized below: 

Photobiological Systems 

• Conduct research to further reduce the truncated antenna of mutants 

• Conduct research to eliminate or alleviate the current Electron Transfer Rate (ETR) 

limitations for organisms and resulting light saturation

• Conduct sensitivity testing on advanced mutants to more fully assess the effects of algae 

photon utilization rate saturation

• Demonstrate the Chemostat II operation at larger scale

• Evaluate alternative reactor bed concepts

• Consider the use of an alternative immobilization mat for use in the B-4 system. As

currently postulated, B-4 uses a fibrous, high surface area polypropylene mat on which 

to cultivate algae. This mat has been shown to work well with other algal organisms and 

can be purchased at low cost as opposed to having to be fabricated on-site (at an 

anticipated high cost). However, alternate mats, perhaps made of alginate, offer several 

advantages: they can be used as a nutrient source for the algae and/or used as feedstock 

for the fermentor when the photobiological system is integrated with an algal fermentor. 

For these reasons, consideration of alternate B-4 systems is warranted. 

• Conduct more comprehensive systems designs, addressing details of mixing, pH control, 

bed temperature control, and CO2 absorption 

Future Algae Fermentation Systems 

• Evaluate feedstock algae pre-treatment processes (e.g. hydrolysis and saccharification) 

to significantly increase the algae-to-hydrogen conversion rate (by greater conversion of

lipids and other components) 

• Consider stand-alone large-scale hydrogen production using high glucose content, rapid 

growing, denser concentration algae as a substrate, grown specifically for fermentation

• Use non-sulfur deprived algae, greatly increasing the easily fermentable glucose content

• Reduce fermentation time from 72 hours to at least the 40 hour batch time characteristic 

of the lignocellulosic fermentation

Lignocellulose Fermentation Systems

• Conduct further analysis to verify the benefit of the lignocellulose fermentation liquid 

organic byproduct sale to dramatically reduce the cost of hydrogen. This would include 

evaluation of potential byproduct component selling price reductions due to increases in 

worldwide supply from fermentation processors 

• Conduct further analysis to address the separation of byproduct components (acids,

ethanol, etc.), as in a distillation process
193 

MEC Systems

• Through research and analysis, develop optimized production process and components 

for low capital cost systems, addressing issues such as:

o Increased solution density 

o Higher pressure operation 

o Low cost cathodes and anodes 

o Cathode and anode optimization for large reactors 

• Determine the extent of ion transport loss increases as reactor size grows to production 

scales

Integrated Systems

• Postulate additional synergistic integrated system combinations to achieve reduced 

hydrogen cost 

• Analyze integration of reactor bed configurations (dual beds, other) that may not be 

optimal for stand-alone photobiological reactors, but are advantageous when combined 

in an integrated system
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4. TITLE AND SUBTITLE

Technoeconomic Boundary Analysis of Biological Pathways to 

Hydrogen Production: March 27, 2008 – August 31, 2009

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DE-AC36-08-GO28308

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Directed Technologies, Inc. 

One Virginia Square

3601 Wilson Boulevard, Suite 650

Arlington, Virginia 22201

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14. ABSTRACT (Maximum 200 Words)

Report documenting the biological and engineering characteristics of five algal and bacterial hydrogen production 

systems selected by DOE and NREL for evaluation.

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technoeconomic analysis; boundary analysis; biological pathways; hydrogen production; Directed Technologies, Inc.

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