A national laboratory of the U.S. Department of Energy 

Office of Energy Efficiency & Renewable Energy 

National Renewable Energy Laboratory 

Innovation for Our Energy Future 

Biomass to Hydrogen Technical Report

NREL/TP-510-37408 Production Detailed Design May 2005 

and Economics Utilizing the 

Battelle Columbus 

Laboratory IndirectlyHeated Gasifier 

P. Spath, A. Aden, T. Eggeman, 

M. Ringer, B. Wallace, and J. Jechura

NREL is operated by Midwest Research Institute ● Battelle Contract No. DE-AC36-99-GO10337 
Biomass to Hydrogen Technical Report

NREL/TP-510-37408 Production Detailed Design May 2005 

and Economics Utilizing the 

Battelle Columbus 

Laboratory IndirectlyHeated Gasifier 

P. Spath, A. Aden, T. Eggeman, 

M. Ringer, B. Wallace, and J. Jechura 

Prepared under Task No(s). BB05.3710 

National Renewable Energy Laboratory 

1617 Cole Boulevard, Golden, Colorado 80401-3393 

303-275-3000 • www.nrel.gov 

Operated for the U.S. Department of Energy 

Office of Energy Efficiency and Renewable Energy 

by Midwest Research Institute • Battelle 

Contract No. DE-AC36-99-GO10337 
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Executive Summary 

The U.S. Department of Energy (DOE) Biomass Program promotes the development of 

technologies for converting biomass into valuable fuels, chemicals, and power that foster 

the growth of biorefineries with the goal of reducing foreign oil imports. With this in 

mind, in 2003, the National Renewable Energy Laboratory (NREL) conducted an 

extensive literature search and examined the technical and economic feasibility of 

numerous fuels and chemicals from biomass-derived syngas (Spath and Dayton, 2003). 

Hydrogen was one product that emerged as highly favorable in this technical and 

economic feasibility study. Therefore, hydrogen was chosen as a model product to 

conduct further analysis and examine the process integration effects and economics of a 

final product from biomass gasification. 

This analysis developed detailed process flow diagrams and an Aspen Plus® model, 

evaluated energy flows including a pinch analysis, obtained process equipment and 

operating costs, and performed an economic evaluation of two process designs based on 

the syngas clean up and conditioning work being performed at NREL. One design, the 

current design, attempts to define today’s state of the technology. The other design, the 

goal design, is a target design that attempts to show the effect of meeting specific 

research goals. Each process design broadly consists of feed handling, drying, 

gasification, gas clean up and conditioning, shift conversion, and purification with some 

unit operation differences. The main difference between the current design and goal 

design is in the tar reformer. The tar reformer in the current design is a bubbling 

fluidized bed reactor with 1% per day catalyst replacement. In the goal design, there is a 

tar reformer/catalyst regenerator system and because the conversion of methane is higher 

for this case, the steam methane reformer can be eliminated from the process design. 

Several parts of the system operate at a high temperature, therefore, heat integration and 

recovery are important. Each process design recovers process heat in a steam cycle with 

an extraction steam turbine/generator to produce some power and supply steam for 

gasification and steam methane reforming or shift conversion. 

Both designs utilize the Battelle Columbus Laboratory (BCL) low pressure indirectlyheated gasifier. The base case plant size is 2,000 dry tonne/day and the feedstock cost is 

$30/dry ton. The current design plant produces 57 MM kg/yr or 66 MM scf/day of 

hydrogen at 100% capacity. The goal design plant produces 61 MM kg/yr or 71 MM 

scf/day of hydrogen at 100% capacity. 

The results of this analysis show a minimum hydrogen selling price of $1.38/kg 

($11.48/GJ, lower heating value [LHV]) for the current design base case analysis and a 

price of $1.24/kg ($10.34/GJ, LHV) for the goal design. The hydrogen price decreases 

for the goal design mainly because of an increase in the hydrogen yield. The decrease in 

the total project investment also has some affect. This result shows that the research at 

NREL in catalytic tar destruction and heteroatom removal is moving in a direction that 

has the potential to decrease the cost of producing clean syngas and any subsequent fuel 

products via biomass gasification. 

iii 
Several sensitivity cases were run to examine the effects of different parameters on the 

analysis. The feedstock cost contributes the most to the product hydrogen price (about 

30%), and thus this variable will always have a large impact on the economics. Overall, 

the sensitivity analysis shows that any parameter that significantly affects the heat 

balance of the system will greatly affect the minimum hydrogen selling price. 

As a benchmark for thermochemical conversion, the DOE Biomass Program is setting 

program targets based on intermediate syngas prices to track progress toward reducing 

the technical barriers associated with biomass gasification. Therefore, this analysis 

included calculations in determining both an intermediate and a stand-alone clean, 

reformed syngas price. The intermediate syngas price for the current and goal designs 

are $6.88/GJ ($7.25/MMBtu) and $4.98/GJ ($5.25/MMBtu), respectively. This is the 

price for clean, reformed syngas as an intermediate in the integrated biomass-to-hydrogen 

design. Stand-alone syngas plants are not being built today, but for a stand-alone plant 

based on the current design, the syngas price would be $8.22/GJ ($8.67/MMBtu), and 

$6.73/GJ ($7.10/MMBtu) for a plant based on the goal design. The lower intermediate 

syngas price shows the importance of integration within the fuels synthesis process plant. 

More detailed capital costs in the feed handling, gasification, and clean up areas would 

improve the accuracy of the analysis. Additionally, more work needs to be done to 

compare indirect gasification with direct gasification to determine the most suitable and 

economically viable gasification system for different fuels products. Future work will 

entail examining other biomass feedstocks and other products along with the integration 

of thermochemical and biochemical conversion processes into biorefinery concepts. 

iv 
Table of Contents 

1.0 Introduction ............................................................................................................................... 1 

2.0 Analysis Approach .................................................................................................................... 1 

3.0 Feedstock and Plant Size .......................................................................................................... 1 

4.0 Process Design Basis ................................................................................................................ 2 

5.0 Current Design Process Overview............................................................................................ 3

6.0 Goal Design Process Overview ................................................................................................ 5 

7.0 Current Design - Process Design, Modeling, and Costing ....................................................... 7 

7.1 Feed Handling and Drying – Area 100 ................................................................................ 7 

7.2 Gasification and Tar Reforming – Area 200 ........................................................................ 7 

7.3 Gas Clean Up and Compression – Area 300...................................................................... 11 

7.4 Reforming, Shift, and PSA – Area 400.............................................................................. 12 

7.5 Hydrogen Compression – Area 500 ................................................................................... 14 

7.6 Steam System and Power Generation – Area 600.............................................................. 14 

7.7 Cooling Water and Other Utilities – Area 700................................................................... 16 

7.8 Additional Design Information .......................................................................................... 17 

8.0 Capital Costs........................................................................................................................... 18 

8.1 Feed Handling, Drying, Gasification and Gas Clean Up Capital Costs............................. 19 

8.2 Other Capital Costs ............................................................................................................ 21 

9.0 Operating Costs ...................................................................................................................... 22 

9.1 Variable Operating Costs ................................................................................................... 22 

9.2 Fixed Operating Costs........................................................................................................ 23 

10.0 Pinch Analysis ...................................................................................................................... 25 

11.0 Energy Balance..................................................................................................................... 27 

12.0 Design, Modeling, and Capital Cost Changes for Goal Design ........................................... 32 

13.0 Resulting Economics of Current Design .............................................................................. 37 

14.0 Current Design Sensitivity Analyses .................................................................................... 38

15.0 Resulting Economics of Goal Design................................................................................... 46 

16.0 Goal Design Sensitivity Analyses ........................................................................................ 48 

17.0 Sensitivity to Plant Size........................................................................................................ 50 

18.0 Syngas Price ......................................................................................................................... 51 

18.1 Intermediate Syngas Price................................................................................................ 51 

18.2 Stand-alone Syngas Price ................................................................................................. 52 

19.0 Hydrogen Program Analysis ................................................................................................ 53 

20.0 Conclusions .......................................................................................................................... 53 

21.0 Future Work ......................................................................................................................... 54 

22.0 References ............................................................................................................................ 55 

Appendix A: Current and Goal Base Case Summary Sheets 

Appendix B: Sensitivity Summary Sheets 

Appendix C: Current Design Process Flow Diagrams 

Appendix D: Goal Design Process Flow Diagrams 

Appendix E: Graphical Correlations for Gas Components and Char 

Appendix F: Flow Charts for Gasifier Elemental Balances 

Appendix G: Equipment Design Parameters and Cost References 

Appendix H: Current Design Summary of Individual Equipment Costs 

Appendix I: Goal Design Summary of Individual Equipment Costs 

v 
List of Figures 

Figure 1: Approach to Process Analysis ........................................................................................ 2 

Figure 2: Block Flow Diagram of Current Design ......................................................................... 4 

Figure 3: Block Flow Diagram of Goal Design ............................................................................. 6 

Figure 4: Current Design Heat Exchange Network within the Steam Cycle ............................... 15 

Figure 5: Current Design Grand Composite Curve ...................................................................... 26 

Figure 6: Current Design Process Energy Balance (LHV Basis) ................................................. 30 

Figure 7: Current Design Process Energy Balance (HHV Basis)................................................. 31 

Figure 8: Goal Design Heat Exchange Network within the Steam Cycle .................................... 34 

Figure 9: Goal Design Grand Composite Curve .......................................................................... 35 

Figure 10: Goal Design Process Energy Balance (LHV Basis) ................................................... 36 

Figure 11: Current Design Base Case Cost Contribution Diagram .............................................. 39 

Figure 12: Current Design Sensitivity Analysis Results .............................................................. 43 

Figure 13: Effect of IRR and Debt/Equity on Current Design Base Case.................................... 44 

Figure 14: Goal Design Base Case Cost Contribution Diagram .................................................. 47 

Figure 15: Goal Design Sensitivity Analysis Results................................................................... 49 

Figure 16: Effect of IRR and Debt/Equity on Goal Design Base Case ........................................ 49 

Figure 17: Effect of Plant Size on Minimum Hydrogen Selling Price ......................................... 51 

Figure 18: Intermediate Syngas Price........................................................................................... 52 

List of Tables 

Table 1: Ultimate Analysis of Hybrid Poplar Feed (wt%, dry basis) ............................................. 1 

Table 2: Tar Reformer Performance - % Conversion to CO & H2................................................. 3 

Table 3: Gasifier Operating Parameters, Yields, and Gas Compositions....................................... 8 

Table 4: Current Design Performance of Tar Reformer ............................................................... 10 

Table 5: Current Design Tar Reformer Properties and Outlet Gas Composition......................... 10 

Table 6: Current Design Plant Power Requirement ..................................................................... 16 

Table 7: Utility and Miscellaneous Design Information .............................................................. 17 

Table 8: Cost Factors in Determining Total Installed Equipment Costs ...................................... 18 

Table 9: Cost Factors for Indirect Costs....................................................................................... 18 

Table 10: Current and Goal Design Base Case TPI Results.......................................................... 19 

Table 11: Feed Handling & Drying and Gasifier & Gas Clean Up Costs from the Literature 

Scaled to 2,000 tonne/day plant............................................................................................ 20 

Table 12: System Design Information for Gasification References ............................................. 20 

Table 13: Variable Operating Costs ............................................................................................. 22 

Table 14: Labor Costs .................................................................................................................. 23 

Table 15: Other Fixed Costs......................................................................................................... 24 

Table 16: Salary Comparison ....................................................................................................... 25 

Table 17: Current Design Overall Energy Analysis (LHV basis) ................................................ 28 

Table 18: Goal Design Performance of Tar Reformer ................................................................. 32 

Table 19: Goal Design Tar Reformer Properties and Outlet Gas Composition ........................... 33 

Table 20: Goal Design Plant Power Requirement........................................................................ 33 

Table 21: Economic Parameters ................................................................................................... 37 

Table 22: Current Design - Sensitivity Analysis Cases................................................................ 40 

Table 23: Current Design - Base Case and Sensitivity Analysis Results ..................................... 42 

Table 24: Goal Design – Sensitivity Analysis Cases ................................................................... 48 

Table 25: Goal Design Base Case and Sensitivity Analysis Results............................................ 50 

Table 26: Stand-alone Syngas Price ............................................................................................. 53 

vi 
Acronyms 

Aspen Plus – Advanced Simulator for Process Engineering 

BCL – Battelle Columbus Laboratory 

BDW – bone dry wood 

BFW – boiler feed water 

DCFROR – discounted cash flow rate of return 

DDB – double declining balance 

DOE – Department of Energy 

EIA – Energy Information Administration 

EOS – equation of state 

GHSV – gas hourly space velocity 

GJ – gigajoule 

gpm – gallons per minute 

GTI – Gas Technology Institute 

HHV – higher heating value 

HTS – high temperature shift 

IP – intermediate pressure 

IRR – internal rate of return 

K.O. – knock out 

kW - kilowatt 

kWh – kilowatt-hour 

LHV – lower heating value 

LP – low pressure 

LTS – low temperature shift 

MACRS – Modified Accelerated Cost Recovery System 

MHSP – minimum hydrogen selling price 

MMBtu – million British thermal units 

MW - megawatt 

NREL – National Renewable Energy Laboratory 

ORNL – Oak Ridge National Laboratory 

PFD – process flow diagram 

PSA – pressure swing adsorption 

Psia – pounds per square inch absolute 

R&D – research and development 

SCF – standard cubic feet 

TIC – total installed cost 

TPEC – total purchased equipment cost 

TPI – total project investment 

VP – vacuum pressure 

vii 
1.0 Introduction 

In 2003, the National Renewable Energy Laboratory performed a preliminary screening 

study of potential products from biomass-derived syngas (Spath and Dayton, 2003). This 

study showed hydrogen to be an economically feasible product, so it was used as a model 

product to show the process integration effects and economics of a final product from 

biomass gasification. In general, the analysis performed for the 2003 study was a highlevel analysis that gathered material and energy balance information along with capital 

and operating cost data from various literature sources. In the case of hydrogen, 

however, NREL had previously developed two Aspen Plus models of hydrogen 

production via gasification. This analysis builds on one of NREL’s models, the indirect 

gasification model. In the original model’s design any excess steam was sold over the 

fence. In the updated model, a steam cycle produces the amount of steam required by the 

plant plus some electricity. Additionally, in this analysis the gas clean up and 

conditioning research work at NREL is also incorporated in the model. 

2.0 Analysis Approach 

The approach that was used in the development of the process designs and economic 

analysis can be seen in Figure 1. For this analysis the first step was to develop process 

flow diagrams (PFDs) and to use these along with literature information and research 

results to build an Aspen Plus model. The energy and material balance from the Aspen 

model were used to size equipment and determine capital and operating costs. 

Additionally, for this analysis, some of the capital costs were obtained from literature 

sources. Once the capital and operating costs are determined, the information is put into 

an excel spreadsheet that is set up to calculate the hydrogen selling price using a 

discounted cash flow rate of return analysis. 

3.0 Feedstock and Plant Size 

The feedstock used for this analysis is hybrid poplar wood chips delivered at 50 wt% 

moisture. The ultimate analysis for the feed used in this study is given in Table 1. The 

plant capacity is designed to be 2,000 bone dry tonne/day. The plant is considered to be 

an “nth” plant design (i.e., established and not a first of a kind or pioneer plant). The 

feedstock cost is assumed to be $30/bone dry ton (delivered) for urban wood waste, 

forest, and mill residues. Information from Oak Ridge National Laboratory (ORNL) 

suggests that the cumulative amount of biomass available at $30/dry ton is 105 million 

tons (http://bioenergy.ornl.gov/resourcedata/index.html). 

Table 1: Ultimate Analysis of Hybrid Poplar Feed (wt%, dry basis) 

Component C H N S O Ash 

wt%, dry basis 50.88 6.04 0.17 0.09 41.90 0.92 

Heating value (Btu/lb): 8,671 HHV 8,060 LHV 

(calculated by Aspen Plususing Boie correlation) 

Source: Craig and Mann (1996) 

1 
( /

) 

Rigorous Material & 

Energy Balance 

ASPEN Plus 

DOE/NREL Sponsored 

Research Results 

gasification gas clean 

up & conditioning

Capital & Project 

Cost Estimation 

Discounted Cash Flow 

Economic Spreadsheet 

ICARUS Software 

Cost Estimation Capital Costs from 

Other Studies 

Hydrogen Selling 

Price 

Process Flow Diagrams 

Literature Information 

Commercial Technology 

Operating Costs 

Fixed & Variable 

Figure 1: Approach to Process Analysis 

4.0 Process Design Basis 

Two process designs were examined in this study. They are based on the current 

operation and performance goals of the catalytic tar destruction and heteroatom removal 

work at NREL. The current design attempts to define today’s state of the technology. 

The goal design is a target design that attempts to show the effect of meeting specific 

research and development (R&D) goals. Table 2 gives the percent conversion of various 

compounds whose concentrations are measured before and after NREL’s tar reformer as 

well as the desired conversion goal (Phillips, et al, 2004). Each process design, both the 

current and goal designs, broadly consists of: 

• feed handling, 

• drying, 

• gasification, 

• gas clean up and conditioning, 

• shift conversion, 

• and hydrogen purification, 

• integrated with a steam and power generation cycle. 

There are some unit operation differences and the details of these two designs will be 

discussed the following sections. 

2
Table 2: Tar Reformer Performance - % Conversion to CO & H2 

Compound Current Design Goal Design 

Methane (CH4) 20% 80% 

Ethane (C2H6) 90% 99% 

Ethylene (C2H4) 50% 90% 

Tars (C10+) 95% 99.9% 

Benzene (C6H6) 70% 99% 

Ammonia (NH3)* 70% 90% 

* Converts to N2 and H2 

5.0 Current Design Process Overview 

A block flow diagram of the current design can be seen in Figure 2. The process flow 

diagrams (PFDs) for this process design are included at the end of this report in Appendix 

C: Current Design Process Flow Diagrams. A more detailed discussion of this process 

can be found in section 7.0 Current Design - Process Design, Modeling, and Costing and 

its subsections. First, the as-received wood is dried from 50 wt% moisture down to 12 

wt% employing a rotary dryer. The dryer uses gas from the char combustor as the drying 

medium. Conveyors and hoppers are used to feed the wood to the low-pressure 

indirectly-heated entrained flow gasifier. Heat for the endothermic gasification reactions 

is supplied by circulating hot synthetic olivine, which is a calcined magnesium silicate 

(primarily Enstatite [MgSiO3] Forsterite [Mg2SiO3], and Hematite [Fe2O3]) used as a 

sand for various applications, between the gasifier and a char combustor vessel. A small 

amount of MgO is added to the fresh olivine to prevent the formation of glass-like bed 

agglomerations that would result from biomass potassium interacting with the silicate 

compounds. The gasification medium is steam. The char that is formed in the gasifier is 

burned in the combustor to reheat the olivine. Particulate removal is performed through 

cyclone separators. Ash and any sand particles that are carried over end up being 

landfilled. 

Gas clean up and conditioning consists of using a tar reformer followed by syngas 

cooling, compression, sulfur removal, steam methane reforming, and high and low 

temperature shift conversion. The tar reformer is a bubbling fluidized bed reactor. 

Catalyst replacement was assumed to be 1% per day of the total catalyst volume (Bain, 

2004). The syngas is cooled through heat exchange with the steam cycle and additional 

cooling via water scrubbing. The scrubber also removes impurities such as particulates 

and ammonia along with any residual tars. The excess scrubber water is sent off site to a 

waste-water treatment facility. The syngas is compressed using a five-section centrifugal 

compressor with interstage cooling. The syngas exiting the gasifier contains almost 400 

ppmv of H2S, therefore sulfur removal is performed using a liquid phase oxidation 

process (LO-CAT®) followed by a ZnO bed. Elemental sulfur is produced and stockpiled 

for disposal. It is stockpiled onsite, instead of being sold or disposed of right away, 

because the amount produced is small and further conditioning would be required before 

the sulfur could be sold. 

3
Steam 

465 psi Cooling Flue gas 23 psi Steam Water 15 psi 

1,598 F cycle 146 F 140 F water 

(63 C) (60 C) Gasification (870 C) 

Tar Scrubber Syngas K.O

Feed Prep Dryer (BCL, & 

Biomass cyclones) reformer Compression vess

19 psi 17 psi 1,383 F 300 F Sand Char & Sand Steam Steam (750 C) (149 C) To treatment cycle cycle Char 

Air 

combustor & cyclones ZnO bed LO-CAT 

Ash 

Steam Steam 

390 psi cycle cycle 405 psi 392 F 

HTS 

1,562 F (200 C) 

445 psi Steam Combustor 707 F 

flue gas (375 C) Steam Sulfur 

Air reformer Steam Natural gas preheat cycle (850 C) Steam 

400 psi Reformer 

662 F (350 C) Cooling LTS Q Combustor Steam F

cycle water 

Off gas K.O. air K.O. H2 vessel cooler Compression vessel 1

370 psi PSA 1110 F 

Process integration heat exchange (43 C) (4

1,265 psi 450 psi 35 psi 1.5 psi

1,000 F 732 F 266 F 115 F 

(538 C) (389 C) (130 C) (46 C) 

IP LP VP Deaerator 

Steam to Steam to Cooling Make up From K.O

reformer gasifier water vessels 

Figure 2: Block Flow Diagram of Current Design 

4 
Reforming (CnHm + nH2O Ù (n+m/2)H2 + nCO) and water-gas shift (CO + H2O Ù CO2 

+ H2) are the main reactions in the steam reformer. The steam reformer is fueled by the 

pressure swing adsorption (PSA) offgas and for burner control a small amount of natural 

gas is added. The high temperature shift (HTS) and low temperature shift (LTS) reactors 

convert the majority of the CO when reacted with H2O into CO2 and H2 through the 

water-gas shift reaction. 

For purification, a pressure swing adsorption unit is used to separate the hydrogen from 

the other components in the shifted gas stream, mainly CO2, and unreacted CO, CH4, and 

other hydrocarbons. For a 70 mol% hydrogen PSA feed, a hydrogen recovery rate of 

85% is typical with a product purity of 99.9 vol%. Finally, the hydrogen is compressed 

to 1,015 psia prior shipment through a pipeline. 

The steam cycle produces power in addition to providing steam for the gasifier and 

reformer operations. The steam cycle is integrated with the biomass-to-hydrogen 

production process. Steam is supplied to the reformer and gasifier from the intermediate 

and low pressure turbine sections of the extraction steam turbine/generator, respectively. 

Superheated steam enters the intermediate pressure turbine at 1,000oF and 1,265 psia and 

is expanded to a pressure of 450 psia. The steam then enters a low pressure turbine and is 

expanded to a pressure of 35 psia. Finally, the steam enters a condensing turbine and is 

expanded to a pressure of 1.5 psia. Preheaters, steam generators, and superheaters are 

integrated within the process design. The condensate from the syngas compressor and 

the condensate from the cooled shifted gas stream prior to the PSA are sent to the steam 

cycle, de-gassed, and combined with the make-up water. A pinch analysis was 

performed to determine the heat integration of the system. 

A cooling water system is also included in the Aspen Plus model to determine the 

requirements of each cooling water heat exchanger within the hydrogen production 

system as well as the requirements of the cooling tower. The cooling water supply 

temperature is 90°F and the return temperature is 110°F. 

6.0 Goal Design Process Overview 

The goal design differs from the current design in that the tar reformer now consists of a 

reactor vessel and a catalyst regeneration vessel. Additionally, since the tar reformer now 

reforms a significant amount of the syngas methane (see Table 2), the steam reformer 

was eliminated from the design. The tar reforming reactor/catalyst regenerator system 

operates isothermally. The heat required for the tar reforming reactor/catalyst regenerator 

system is supplied by burning the PSA offgas along with some natural gas. The steam to 

carbon ratio for the shift conversion step is set at 2 mol of H2O/mol of C. The biomassto-hydrogen process is integrated with the steam cycle. A block flow diagram of the goal 

design is shown in Figure 3. Additionally, process flow diagrams (PFDs) for this process 

design are included at the end of this report in Appendix D: Goal Design Process Flow 

Diagrams and more detailed information can be found in section 12.0 Design, Modeling, 

and Capital Cost Changes for Goal Design. 

5
Steam 

Flue gas 23 psi Steam Water 15 psi 

1,598 F cycle 146 F 

(63 C) 

Tar 

Gasification (870 C) 

Scrubber Syngas Feed Prep Dryer (BCL, & cyclones) reformer 19 psi Biomass compressi

1,598 F 17 psi To treatment Sand Char & Sand 156 F 

Catalyst (870 C) (69 C) Steam Char Natural gas cycle combustor & Tar reformer Air cyclones Air catalyst regenerator Flue gas 

Ash 

Steam Steam Cooling Steam 445 psi Steam 390 psi cycle 707 F cycle cycle water

392 F (375 C) (200 C) K.O. HTS ZnO bed LO-CAT vessel 

400 psi 455 psi 

686 F LTS Steam (363 C)

cycle

Sulfur 120 F 

Cooling (49 C) 

water Off gas 

K.O. H

air K.O. PSA Compression vessel cooler vessel 1

370 psi 110 F 1(43 C) (4

i

IP LP 

Cooling 

(46 C) 

711 F 

Steam to Make up shift 

Steam to 

fier 

From vesse

VP Deaerator 

Process integration heat exchange 

1.5 psi 115 F 

35 psi 266 F (130 C) 

410 psi (377 C) 

1,265 psi 1,000 F (538 C) 

gas water 

Figure 3: Block Flow Diagram of Goal Design 

6 
7.0 Current Design - Process Design, Modeling, and Costing 

The following sections describe the detailed process design for the current design as 

outlined in section 5.0 Current Design Process Overview. 

7.1 Feed Handling and Drying – Area 100 

The feed handling and drying section are shown in PFD-P700-A101 and PFD-P700-

A102. Wood chips are delivered to the plant primarily via trucks. However, it is 

envisioned that there could be some train transport. Assuming that each truck capacity is 

about 25 tons (Mann and Spath, 1997), this means that if the wood, at a moisture content 

of 50%, was delivered to the plant via truck transport only, then 176 truck deliveries per 

day would be required. As the trucks enter the plant they are weighed (M-101) and the 

wood chips are dumped into a storage pile. From the storage pile, the wood chips are 

conveyed (C-102) through a magnetic separator (S-101) and screened (S-102). Particles 

larger than 2 inches are sent through a hammer mill (T-102/M-102) for further size 

reduction. Front end loaders transfer the wood chips to the dryer feed bins (T-103). 

Because of the large plant size there are two identical, parallel feed handling and drying 

trains. The wet wood chips enter each rotary biomass dryer (M-104) through a dryer feed 

screw conveyor (C-104). After drying the wood to a moisture content of 12 wt% with 

flue gas from the char combustor (R-202), the gas is sent through a cyclone (S-103) and 

baghouse filter (S-104) to remove any particulates prior to being emitted to the 

atmosphere. The stack temperature of the flue gas is set at 250°F, which is above the 

dew point of the gas. The stack temperature is controlled by cooling the hot flue gas (H101) prior to entering the dryer. This heat is used to generate steam (see section 7.6 

Steam System and Power Generation – Area 600). The dried biomass is then conveyed 

to the gasifier train (T-104/C-105). 

7.2 Gasification and Tar Reforming – Area 200 

From the feed handling and drying section, the dried wood enters the gasifier and tar 

reforming section as shown in PFD-P700-A201 and PFD-P700-A202. Because of the 

plant size, it is assumed that there are two gasifier trains. The gasifier (R-201) used in 

this analysis is a low-pressure indirectly-heated entrained flow gasifier. 

The gasifier was modeled using correlations based on run data from the Battelle 

Columbus Laboratory (BCL) 9 tonne/day test facility. The data and correlations for the 

gasifier can be found in Bain (1992). The experimental runs were performed for several 

different wood types including Red Oak chips, Birch and Maple chips, Pine chips, 

sawdust, and other hard and soft wood chips. The original pilot plant data for these runs 

can be found in Feldmann, et al, (1988). The temperature range for the data is 1,280-

1,857oF and the pressure range is 2.4-14.4 psig with the majority of the data being in the 

1,500-1,672oF range. 

7
--- 

The BCL test facility data was regressed using a polynomial function (Bain, 1992). The 

quadratic function (A + B*T + C*T2

) provides a good fit for the conversion of all of the 

gas components and the char. The correlations are in terms of standard cubic feet (scf) of 

component/lb of BDW except for the char and tar, which are in terms of lb of 

component/lb of BDW. Graphs of the correlations can be found in Appendix E: 

Graphical Correlations for Gas Components and Char. These correlations along with 

documentation have been programmed into a Fortran file. Aspen Plus passes the gasifier 

temperature to the Fortran file, the Fortran file uses the correlations to calculate the gas 

and char yields then elemental balances are performed for carbon, hydrogen, sulfur, and 

oxygen to come up with the overall material and energy balance for the gasifer. The 

elemental balances were put into flow charts and are included in Appendix F: Flow 

Charts for Gasifier Elemental Balances. The BCL model performs the elemental 

balances in the following order: carbon, oxygen, sulfur, and hydrogen. However, in 

general, the sulfur balance can be performed any time as long as it is done before the 

hydrogen balance. Note, when running the Aspen Plus model it is important for the user 

to look at the history file for errors, make any necessary changes, re-run the model, and 

examine the history file again when changing any of the model parameters. 

Table 3 gives the resulting operating parameters, tar and char yields, and gas composition 

for the BCL gasifier from the Aspen Plus model. 

Table 3: Gasifier Operating Parameters, Yields, and Gas Compositions 

Gasifier Variable Value 

Gasifier type BCL 

Temperature 1,598oF (870oC) 

Pressure 23 psia (1.6 bar) 

Steam/bone dry feed 0.4 lb/lb 

Sand purge 0.1wt% of circulation rate 

Gas composition mol% (wet) mol% (dry) 

H2 12.91 23.85 

CO2 6.93 12.79 

CO 22.84 42.18 

H2O 45.87 

CH4 8.32 15.36 

C2H2 0.22 0.41 

C2H4 2.35 4.35 

C2H6 0.16 0.29 

C6H6 0.07 0.13 

tar (C10H8) 0.13 0.23 

NH3 0.18 0.32 

H2S 0.04 0.07 

Gas yield 0.04 lb-mol of dry gas/lb bone dry feed 

Gas heating value (Btu/lb) Wet: 4,739 HHV 4,402 LHV 

Dry: 7,984 HHV 7,417 LHV 

Char yield 0.22 lb/lb bone dry feed 

H2:CO molar ratio 0.57 

Gasifier efficiency 72.1% HHV basis 

71.8% LHV basis 

8 
Note: The gasifier efficiency is defined as the combustion energy of the synthesis gas divided by 

the combustion energy of the biomass. 

Heat for the endothermic gasification reactions is supplied by circulating a hot medium 

between the gasifier vessel and the char combustor. In this case the medium is synthetic 

olivine, a calcined magnesium silicate (primarily Enstatite [MgSiO3] Forsterite 

[Mg2SiO3], and Hematite [Fe2O3]) used as a sand for various applications. 

A small amount of MgO must be added to the fresh olivine to avoid the formation of 

glass-like bed agglomerations that would result from the biomass potassium interacting 

with the silicate compounds. The MgO titrates the potassium in the feed ash. Without 

MgO addition, the potassium will form glass, K2SiO4, with the silica in the system. 

K2SiO4 has a low melting point (~930°F) and its formation will cause the bed media to 

become sticky, agglomerate, and eventually defluidize. Adding MgO makes the 

potassium form a high melting (~2,370°F) ternary eutectic with the silica, thus 

sequestering it. Potassium carry over in the gasifier/combustor cyclones is also 

significantly reduced. The ash content of the feed is assumed to contain 0.2 wt% 

potassium. The MgO flow rate is set at two times the molar flow rate of potassium. 

The gasification medium is steam which is supplied from the steam cycle (7.6 Steam 

System and Power Generation – Area 600). The steam-to-wood ratio is 0.4 lb of steam/lb 

of bone dry wood. This variable was tested in the sensitivity analysis. The char 

combustor temperature is set at 1,800°F and the gasifier temperature is obtained from the 

energy balance around the gasifier and combustor. The resulting gasifier temperature is 

1,598oF. The gasifier pressure is 23 psia. The olivine circulating flow rate is 27 lb of 

olivine/lb of bone dry wood. Fresh olivine is made up at a rate of 0.11% of the 

circulating rate to account for the losses from the cyclones. The combustion air is varied 

from 5%-20% excess air until the heat duty of the char combustor is zero. The resulting 

excess air turns out to be 12%. 

Particulate removal is performed through cyclone separators. The majority of the olivine 

and char (99.9% of both) is separated in the primary gasifier cyclone (S-201) and sent to 

the char combustor. A secondary cyclone (S-202) removes 90% of any residual fines. 

The char that is formed in the gasifier is burned in the combustor to reheat the olivine. 

The primary combustor cyclone (S-203) separates the olivine (99.9%) from the 

combustion gases and the olivine is sent back to the gasifier. Ash and any sand particles 

that are carried over are removed in the secondary combustor cyclone (99.9% separation 

in S-204) followed by an electrostatic precipitator (S-205) which removes the remaining 

residual amount of solid particles. The sand and ash mixture is landfilled but prior to this 

the solids are cooled and then water is added to the sand/ash stream for conditioning to 

prevent the mixture from being too dusty to handle. First the ash and sand mixture is 

cooled to 300°F using the water cooled screw conveyor (M-201) then water is added 

directly to the mixture until the mixture water content is 10 wt%. 

The gas from the secondary gasifier cyclone is sent to the tar reformer (R-203). In this 

bubbling fluidized bed reactor the following compounds are converted to CO and H2: 

CH4, C2H6, C2H4, C6H6, and C10+; while NH3 is converted to N2 and H2. In the 

9
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simulation, the percent conversion of each compound is set by the conversion amount 

that is currently seen in the catalytic tar destruction and heteroatom removal work at 

NREL. Table 4 gives the conversion that has been experimentally verified from the data 

gathered at NREL’s bench-scale thermo-catalytic conversion system and NREL’s 

Thermochemical Pilot Process Development Unit (TCPDU) (Phillips, et al, 2004). 

Table 4: Current Design Performance of Tar Reformer 

Compound Percent Conversion 

to CO & H2 

Methane (CH4) 20% 

Ethane (C2H6) 90% 

Ethylene (C2H4) 50% 

Tars (C10+) 95% 

Benzene (C6H6) 70% 

Ammonia (NH3)* 70% 

* Converts to N2 and H2 

In the Aspen Plus simulation, the gas entering the tar reformer is at the gasifier 

temperature (1,598oF) and the gas exiting the tar reformer ends up at 1,383°F. The 

composition of the gas from the tar reformer can be seen in Table 5. Prior to the quench 

step, the hot gas is cooled to 300°F with heat exchange (H-201 and H-202) that is 

integrated in the steam cycle (see section 7.6 Steam System and Power Generation – 

Area 600). 

Table 5: Current Design Tar Reformer Properties and Outlet Gas Composition 

Tar Reformer Variable Value 

Tar reformer inlet temperature 1,598oF (870oC) 

Tar reformer outlet temperature 1,383oF (750oC) 

Tar reformer outlet gas composition mol% (wet) mol% (dry) 

H2 33.44 45.52 

CO2 16.10 21.92 

CO 16.51 22.47 

H2O 26.54 

CH4 6.06 8.25 

C2H2 0.10 0.14 

C2H4 1.07 1.46 

C2H6 0.01 0.02 

C6H6 0.02 0.03 

tar (C10H8) 0.01 0.01 

NH3 0.05 0.07 

H2S 0.04 0.05 

N2 0.06 0.08 

Gas heating value (Btu/lb) Wet: 4,979 HHV 4,485 LHV 

Dry: 6,711 HHV 6,045 LHV 

H2:CO molar ratio 2.03 

10 
7.3 Gas Clean Up and Compression – Area 300 

After direct cooling of the syngas to a temperature of 300°F additional cooling is carried 

out via water scrubbing, shown in PFD-P700-A301. The scrubber also removes 

impurities such as particulates, ammonia, and any residual tars. The scrubbing system 

consists of a venturi scrubber (M-302) and quench chamber (M-301). The quench water 

is cooled and recirculated. The quench water flow rate is determined by adjusting the 

circulation rate until the exit temperature from the cooling water heat exchanger (H-301) 

is 110°F. The excess scrubber water is sent off site to a waste water treatment facility. 

This amounts to about 2 gallons per minute excess water for the 2,000 bone dry 

tonne/day plant. Any solids that settle out in T-301 are sent off-site for treatment as well. 

For modeling purposes, the water content of the sludge stream was set at 50 wt%. The 

quench step cools the syngas to a temperature of 140°F. The syngas is then compressed 

using a five-section centrifugal compressor with interstage cooling as shown in PFDP700-A302 (K-301A/B/C/D/E, S-301, S-302A/B/C/D/E, S-303, H-302A/B/C/D/E, and 

H-303). The compressor was modeled such that each section has a polytropic efficiency 

of 78% along with intercooler temperatures of 140°F. 

Sulfur compounds are the main poison of reforming catalysts. Low temperature shift 

catalysts are also very sensitive to sulfur. Because the syngas exiting the gasifier 

contains almost 400 ppmv of H2S, a ZnO bed by itself could not be used for sulfur 

removal. The normal sulfur concentration at the inlet of a ZnO bed is typically 10-20 

ppmv H2S. The ZnO bed will then reduce the sulfur to less than 1 ppmv H2S. A very 

low concentration of less than 1 ppmv H2S is required for steam reforming and the LTS 

catalyst. Even at a concentration of 0.1 ppm the reforming catalyst can start to 

deactivate. Therefore, sulfur removal via a liquid phase oxidation process followed by a 

ZnO bed was chosen. PFD-P700-A303 shows the sulfur removal step. The LO-CAT 

process will remove the bulk of the sulfur but it cannot reliably reduce the sulfur 

concentration to the extremely low levels required by the downstream conversion steps. 

Therefore, two ZnO beds in series follow the LO-CAT process because the H2S 

requirement is so low and a ZnO bed is a simple, relatively inexpensive piece of 

equipment with a known history for reducing H2S concentrations to very low levels. 

Additionally, each ZnO reactor contains a layer of hydrogenation catalyst to convert 

organic sulfur to H2S because it is possible that mercaptans, COS, and other sulfur 

compounds could be present in very small amounts in the syngas stream from the 

gasifier. 

Although, there are several liquid phase oxidation processes available today, the LOCAT process was selected because of its progress in minimizing catalyst degradation and 

its environmentally benign catalyst compared to others. LO-CAT is an iron chelate based 

process that consists of a venturi (M-303), absorber (M-304), oxidizer (R-301), air 

blower (K-302), solution circulation pump (P-303) and solution cooler (H-305). 

Elemental sulfur is produced and since there is such a small amount (1.6 tonne/day), it is 

stockpiled for eventual disposal rather than conditioned and sold. The LO-CAT process 

was modeled to remove the sulfur to a concentration of 10 ppm H2S and the ZnO bed 

removes the remaining sulfur to a concentration of less than 1 ppm. The air flow rate for 

11
re-oxidizing the LO-CAT solution was included in the simulation and calculated based on 

the requirement of 2 mol of O2 per mol of H2S. Prior to entering the LO-CAT system the 

gas stream is superheated to 10°F above dew point (H-304) which in this process is 

equivalent to 120 ̊F. This degree of superheating is required for the LO-CAT system. 

The ZnO bed operates at higher temperatures which are needed so that the reaction (ZnO 

+ H2S Ù ZnS + H2O) closely approaches equilibrium. Therefore, the gas stream exiting 

the LO-CAT process is heated to 707°F (H-306) using heat from the steam cycle (see 

section 7.6 Steam System and Power Generation – Area 600) prior to entering the ZnO 

reactors (R-302). During operation over a length of time, the reaction zone will gradually 

move down through the ZnO bed until the bed material finally needs to be changed out. 

7.4 Reforming, Shift, and PSA – Area 400 

There is a significant amount of CO, CH4 and other hydrocarbons in the biomass derived 

syngas (as can be seen in Table 3), thus these components need to undergo conversion via 

reforming (CnHm + nH2O Ù (n+m/2)H2 + nCO) and shift conversion (CO + H2O Ù CO2 

+ H2) reactions. The steam reformer is shown in PFD-P700-A401. Reforming and 

water-gas shift are the main reactions in the steam reformer. The reforming reaction is 

highly endothermic and is favored by high temperatures and low pressures. The shift 

reaction is exothermic and favors low temperatures and higher steam ratios. The steam 

reformer (R-401) is comprised of catalyst-filled tubes, surrounded by a firebox that 

provides the heat necessary for the endothermic reforming reaction. The main 

components of the reformer furnace include an air/fuel combustion system, a radiant heat 

transfer section, and a convection section. The radiant section supplies heat to the 

catalyst tubes by combusting the air/fuel mixture and the convection section recovers heat 

by cooling down the flue gases (H-401 and H-404). Reformer furnaces are not very 

efficient and only about half of the heat in the radiant section is absorbed by the furnace 

tubes. Generally, the feed gas flows up through the catalyst tubes but the reformer 

furnace can be side-, terrace-, top-, or bottom-fired (Spath and Dayton, 2003). 

Steam reformers typically operate at 1,500-1,600°F and between 218-435 psia using a 

nickel based catalyst. In this analysis the steam reformer was simulated as an equilibrium 

reactor at 1,562°F with a -20°F approach temperature, an inlet pressure of 435 psia, and a 

steam to carbon ratio of 3 mol of H2O/mol of C (Leiby, 1994). The approach temperature 

is defined as the difference between the measured outlet temperature and the temperature 

that would yield the measured conversion of a component at equilibrium (In this case the 

component is methane.). In Aspen Plus the sign of the approach temperature for this 

conversion step is negative but other software packages may use a different convention. 

In this instance, a positive sign would be erroneous resulting in a methane conversion 

which is higher than that obtained at equilibrium. The steam for the reformer is supplied 

from the steam cycle (see section 7.6 Steam System and Power Generation – Area 600). 

The pressure drop through the steam reformer is 30 psi. The reformer is fueled by the 

PSA offgas and a small amount of natural gas is added for burner control. The amount of 

natural gas that is added is equal to 10% of the heating value of the PSA offgas. 

Following the steam reformer, the HTS and LTS reactors convert the majority of the 

remaining CO, when reacted with H2O, into CO2 and H2 through the water-gas shift 

12
reaction. PFD-P700-A402 depicts these shift reactors. The gas exiting the reformer is 

first cooled to 662°F (H-402) (the operating range of a HTS reactor is typically 570-

840°F). The HTS (R-402) and LTS (R-403) were modeled as fixed bed equilibrium 

reactors with approach temperatures of 35°F and 20°F, respectively, (Leiby, 1994). In 

this case for the shift conversion reaction the sign convention for the approach 

temperature in Aspen Plus is positive. In this instance, a negative number would result in 

more CO being converted than is possible at equilibrium. The gas exiting the HTS 

reactor is cooled to 392°F (H-405 and H-406) prior to entering the LTS reactor (The LTS 

reactor typically operates between 350-515°F and often operates near condensation 

conditions.). The HTS catalyst has an iron oxide, chromium oxide basis while the major 

component in the LTS catalyst is copper oxide, most often in a mixture with zinc oxide 

(Spath and Dayton, 2003). 

For purification, a PSA unit is used to separate the hydrogen from the other components 

in the shifted gas stream, mainly CO2, and unreacted CO, CH4, and other hydrocarbons. 

The PSA unit can be seen in PFD-P700-A403. The hydrogen purity achieved from a 

PSA unit can be greater than 99.99+%. Based on past conversations with industrial gas 

producers, the shifted gas stream must contain at least 70 mol% hydrogen before it can be 

economically purified in the PSA unit (Mann, 1995). Purification of streams more dilute 

than this decreases the product purity and recovery of hydrogen. For this analysis, the 

concentration of hydrogen in the shifted stream prior to the PSA is between 60-65 mol%. 

Therefore, part of the PSA hydrogen product stream is recycled back into the PSA feed. 

For a 70 mol% hydrogen PSA feed, a hydrogen recovery rate of 85% is typical with a 

product purity of 99.9 vol%. Prior to the PSA unit, entrained liquids (water and 

condensed hydrocarbons) must be removed because they will permanently damage the 

adsorbent, which is a mixture of activated carbon and zeolites. Cooling the product and 

installing a knock out drum with a mist eliminator (S-401 and S-402) prior to the PSA 

unit is usually sufficient. The PSA efficiency is also affected by adsorption temperature. 

Fewer impurities are adsorbed at higher temperatures because the equilibrium capacity of 

the molecular sieves decreases with increasing temperature. Therefore, the design for 

this analysis uses a heat exchanger integrated with the steam cycle (see section 7.6 Steam 

System and Power Generation – Area 600) to cool the gas down to its dew point (H-407). 

The stream is further cooled by an air-cooled heat exchanger (H-408) to 140°F. A 

cooling water heat exchanger (H-409) is then used to reduce the stream temperature to 

110°F. 

The minimum pressure ratio between the feed and purge gas of the PSA unit is about 4:1. 

The absolute pressures of the feed and purge gas are also important in regard to hydrogen 

recovery. The optimum feed pressure for refinery applications is in the range of 215-415 

psia. The purge gas pressure is typically between 17-20 psia to obtain a high recovery of 

hydrogen. Hydrogen recovery is usually 85-90% at these conditions and drops to 60­

80% at high purge gas pressures of 55-95 psia (Leiby, 1994). In the design for this 

analysis the pressure of the PSA feed gas is 360 psia and the purge gas pressure is 20 

psia. 

13
7.5 Hydrogen Compression – Area 500 

Ultimately, the hydrogen is sent to a pipeline so the product hydrogen is compressed 

from 360 psia to 1,015 psia. This is done using a two-stage reciprocating compressor 

with an isentropic efficiency of 82% and interstage intercooler temperatures of 140°F 

each (K-501A/B, H-501A/B, S-502, H-502, and S-503). PFD-P700-A500 shows the 

hydrogen compression step. 

7.6 Steam System and Power Generation – Area 600 

The process design includes a steam cycle that produces steam via heat recovery of the 

hot process streams throughout the plant. Because the gasifier and reformer both require 

steam, power is produced from the steam cycle using an extraction steam 

turbine/generator (M-602). Steam is supplied to the reformer from the intermediate 

pressure turbine stage and to the gasifer from the low pressure turbine stage. The steam 

system and power generation area is shown in PFD-P700-A601, -A602, and -A603. 

A condensate collection tank (T-601) gathers condensate from the syngas compressor and 

from the cooled shifted gas stream prior to the PSA along with the steam turbine 

condensate and make-up water. The total condensate stream is heated to the saturation 

temperature and sent to the deaerator (T-603) to de-gas any dissolved gases out of the 

water. The water from the deaerator is first pumped to a pressure of 1,345 psia and then 

pre-heated to the saturation temperature using a series of exchangers. The saturated 

steam is collected in the steam drum (T-604). To prevent solids build up, water must be 

periodically discharged from the steam drum. The blowdown rate is equal to 2% of water 

circulation rate. The saturated steam from the steam drum is superheated with another 

series of exchangers. The superheated steam temperature and pressure were set based on 

standard conditions given in Perry, et al, 1997. Superheated steam enters the 

intermediate pressure turbine stage at 1,000oF and1,265 psia and is expanded to a 

pressure of 450 psia where a slipstream is removed and sent to the steam methane 

reformer. The remaining steam then enters the low pressure turbine and is expanded to a 

pressure of 35 psia. Here a slipstream of steam is removed and sent to the gasifier. 

Finally, the steam enters a condensing turbine and is expanded to a pressure of 1.5 psia. 

The steam is condensed in the steam turbine condenser (H-601) and re-circulated back to 

the condensate collection tank. 

A pinch analysis was performed to determine the heat integration of the system (see 

section 10.0 Pinch Analysis for details). Heat integration is an important part of this 

thermal conversion process. Figure 4 is a drawing that shows the heat exchange network 

within the steam cycle. The heat duty of the various sections and the heat exchanger tag 

numbers are given. The figure shows where heat is exchanged between the different 

steps within the process and the steam cycle but it does not show the integration of the 

individual heat exchangers. The integration can be seen on the PFDs (Appendix C: 

Current Design Process Flow Diagrams). In order to close the heat balance of the 

14
Figure 4: Current Design Heat Exchange Network within the Steam Cycle 

15 
system, the Aspen Plus model increases or decreases the water flow rate through the 

steam cycle until the heat balance of the system is met. 

The analysis assumes that all drives for compressors, pumps, fans, etc are electric motors. 

Additionally, 10% excess power is added to total power requirement to account for 

miscellaneous usage. Table 6 contains the power requirement of the plant broken out 

into the different plant sections. Syngas compression accounts for the largest power use. 

Even though the plant produces power, it is not enough to meet the total electricity 

demand of the plant. Therefore, the shortage is made up from electricity that is 

purchased from the grid. 

Table 6: Current Design Plant Power Requirement 

Plant Section Power Requirement (kW) 

Feed handling & drying 742 

Gasification, Tar reforming, & quench 3,636 

Compression & sulfur removal 21,871 

Steam methane reforming, shift, and PSA 630 

Hydrogen compression 3,899 

Steam system & power generation 660 required 

25,583 generated 

Cooling water & other utilities 1,110 

Miscellaneous 3,255 

Total plant power requirement 35,803 

Grid electricity requirement 10,219 

7.7 Cooling Water and Other Utilities – Area 700 

The cooling water system is shown on PFD-P700-A701. A mechanical draft cooling 

tower (M-701) provides cooling water to several heat exchangers in the plant. The tower 

utilizes large fans to force air through circulated water. Heat is transferred from the water 

to the surrounding air by the transfer of sensible and latent heat. Cooling water is used in 

the following pieces of equipment: 

•	 the sand/ash cooler (M-201) which cools the sand/ash mixture from the 

gasifier/combustor 

•	 the quench water recirculation cooler (H-301) which cools the water used in the 

syngas quench step 

•	 the water-cooled aftercooler (H-303) which follows the syngas compressor and 

cools the syngas after the last stage of compression 

•	 the LO-CAT absorbent solution cooler (H-305) which cools the solution that 

circulates between the oxidizer and absorber vessels 

•	 the PSA water-cooled precooler (H-409) which cools the gas in order to condense 

out any liquids prior to the PSA unit 

•	 the hydrogen compressor water-cooled aftercooler (H-502) which follows the 

hydrogen compressor and cools the hydrogen after the last stage of compression 

•	 the blowdown water-cooled cooler (H-603) which cools the blowdown stream 

16
•	 the steam turbine condenser (H-601) which condenses the steam exiting the steam 

turbine 

Make-up water for the cooling tower is supplied at 14.7 psia and 60°F. Water losses 

include evaporation, drift which is the water entrained in the discharge vapor, and 

blowdown. Drift losses were estimated to be 0.2% of the water supply. Evaporation 

losses and blowdown were calculated based on information and equations in Perry, et al, 

1997. The cooling water supply pressure is 65 psia and the supply temperature is 90°F. 

The cooling water return temperature is 110°F. 

An instrument air system is included to provide compressed air for both service and 

instruments. The instrument air system is shown on PFD-P700-A701. The system 

consists of an air compressor (K-701), dryer (S-701) and receiver (T-701). The 

instrument air is at a pressure of 115 psia, a dew point of -40°F, and oil free. 

Other miscellaneous items that are taken into account in the design include: 

•	 a firewater storage tank (T-702) and pump (P-702) 

•	 a diesel tank (T-703) and pump (P-703) to fuel the front loaders 

•	 an olivine truck scale with dump (M-702) and an olivine lock hopper (T-705) as 

well as an MgO lock hopper (T-706) 

• a hydrazine storage tank (T-707) and pump (P-705) 

This equipment is shown on PFD-P700-A702. 

7.8 Additional Design Information 

Table 7 contains some additional information used in the Aspen Plus model and biomass 

gasification to hydrogen production design. 

Table 7: Utility and Miscellaneous Design Information 

Item Design Information 

Ambient air conditions (1,2, and 3) Pressure: 14.7 psia 

TDry Bulb: 90°F 

TWet Bulb: 80°F 

Composition (mol%): 

N2: 75.7% O2: 20.3% Ar: 0.9% CO2: 0.03% H2O: 3.1% 

Pressure drop allowance Syngas compressor intercoolers = 2 psi 

Heat exchangers and packed beds = 5 psi 

Thermodynamics - VLE: Redlich-Kwong-Soave EOS with Boston-Mathias modification. 

- Enthalpies for Non-conventional components: Boie correlation for heat 

of combustion, Kirov correlation for heat capacity. 

- Steam System: ASME Steam Tables. 

(1) In Gas Processors Suppliers Association (2004), see Table 11.4 for typical design 

values for dry bulb and wet bulb temperature by geography. Selected values would cover 

summertime conditions for most of lower 48 states. 

(2) In Weast (1981), see F-172 for composition of dry air. Nitrogen value adjusted 

slightly to force mole fraction closure using only N2, O2, Ar, and CO2 as air components. 

17 
(3) In Perry, et al, (1997), see psychrometric chart, Figure 12-2, for moisture content of 

air. 

8.0 Capital Costs 

The following sections discuss the methods and sources for determining the capital cost 

of each piece of equipment within the plant. A summary of the individual equipment 

costs for the current design can be found in Appendix H: Current Design Summary of 

Individual Equipment Costs and a summary of the individual equipment costs for the 

goal design can be found in Appendix I: Goal Design Summary of Individual Equipment 

Costs. 

Because the majority of the costs came from literature and Questimate® (an equipment 

capital cost estimating software tool by Aspen Tech) instead of vendor quotes, the 

purchased cost of the equipment was calculated and then cost factors were used to 

determine the installed equipment cost. The cost multipliers were taken from Peters and 

Timmerhaus, 2003. This method of cost estimation has an expected accuracy of roughly 

+ or -30%. The factors used in determining the total installed cost (TIC) of each piece of 

equipment are shown in Table 8. 

Table 8: Cost Factors in Determining Total Installed Equipment Costs 

% of TPEC 

Total Purchased Equipment Cost (TPEC) 100 

 Purchased equipment installation 39 

 Instrumentation and controls 26 

 Piping 31 

 Electrical systems 10 

Buildings (including services) 29 

 Yard improvements 12 

Total Installed Cost (TIC) 247 

The indirect costs which are the nonmanufacturing fixed-capital investment costs also 

need to be calculated. These costs were also determined using cost factors taken from 

Peters and Timmerhaus, 2003. The factors are shown in Table 9 and have been put as 

percentages in terms of total purchased equipment cost, total installed cost, and total 

project investment. The total project investment (TPI) is the sum of the total installed 

cost (TIC) plus the total indirect costs. 

Table 9: Cost Factors for Indirect Costs 

Indirect Costs % of TPEC % of TIC % of TPI 

 Engineering 32 13 9 

 Construction 34 14 9 

Legal and contractors fees 23 9 6 

 Project contingency 37 15 10 

Total Indirect Costs 126 51 34 

18 
Table 10 gives the TPI results for the base case 2,000 tonne/day plant current and goal 

case designs. To see the detailed capital costs refer to Appendix H: Current Design 

Summary of Individual Equipment Costs and Appendix I: Goal Design Summary of 

Individual Equipment Costs. 

Table 10: Current and Goal Design Base Case TPI Results 

Cost 2002 $MM 

Current 

Design 

Goal 

Design 

Total Purchased Equipment Cost (TPEC) 41 39 

 Purchased equipment installation 16 15 

 Instrumentation and controls 11 10 

 Piping 13 12 

 Electrical systems 4 4 

Buildings (including services) 12 11 

 Yard improvements 5 5 

Total Installed Cost (TIC) 102 96 

Indirect Costs 

 Engineering 13 12 

 Construction 14 13 

Legal and contractors fees 9 9 

 Project contingency 14 14 

Total Indirect Costs 52 49 

Total Project Investment (TPI) 154 144 

8.1 Feed Handling, Drying, Gasification and Gas Clean Up Capital Costs 

The biomass handling and drying costs as well as the gasification and gas clean up costs 

were obtained from several reports by others that documented detailed design and cost 

estimates. Some of the reports gave costs for individual pieces of equipment while others 

lumped the equipment costs into areas. The costs from the reports were amalgamated 

into (1) feedstock handling and drying and (2) gasification and clean up. Costs from 

those reports scaled to a 2,000 bone dry tonne/day plant are given in Table 11. The costs 

are divided into two types of systems: (1) a low pressure indirectly heated gasifier system 

using the BCL gasifier and (2) a high pressure directly heated gasifier system using the 

Gas Technology Institute (GTI). Table 12 gives the basic dryer and gasifier design basis 

for the references. The base case in this analysis uses the average feed handling and 

drying cost from all of the literature sources and the average gasifier and gas clean up 

cost for the references using the BCL gasifier. A sensitivity analysis was performed to 

examine the effects of these varying study costs. 

19
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Table 11: Feed Handling & Drying and Gasifier & Gas Clean Up Costs from the 

Literature Scaled to 2,000 tonne/day plant 

Reference Scaled Feed Handling 

and Drying Cost $K 

(2002) 

BCL - Scaled 

Gasifier and Gas 

Clean Up Cost $K 

(2002) 

GTI - Scaled 

Gasifier and 

Gas Clean Up 

Cost $K (2002) 

Breault and Morgan (1992) (a) $15,048 $15,801 

Dravo Engineering Companies 

(1987) (a) 

$14,848 $15,774 

Weyerhaeuser, et al, (2000) (a) $21,241 $24,063 

Stone & Webster, et al, (1995) (a) $25,067 $36,232 

Wan and Malcolm (1990) (a) $18,947 (b) 

$14,098 (c) 

$11,289 (b) 

$11,109 (c) 

Weyerhaeuser (1992) (a) $13,468 $10,224 

Wright and Feinberg (1993) (a) $26,048 – BCL design 

$21,942 – GTI design 

$12,318 - quench (d) 

$26,562 - HGCU (d) 

$38,605 

Craig (1994) $13,680 $48,229 

AVERAGE $18,840 $16,392 $41,071 

(a) From detailed design and cost estimates 

(b) Estimated from a 200 dry ton/day plant design. 

(c) Estimated from a 1,000 dry ton/day plant design. 

(d) Two separate gas clean up configurations were examined for the BCL gasifier. 

HGCU = hot gas clean up. 

Table 12: System Design Information for Gasification References 

Reference Feed Handling and BCL Gasifier and GTI Gasifier and 

Drying Gas Clean Up Gas Clean Up 

Breault and Morgan (1992) Rotary dryer Cyclones, heat 

exchange & scrubber 

Dravo Engineering Companies Rotary drum dryer Cyclones, heat 

(1987) exchange & scrubber 

Weyerhaeuser, et al, (2000) Steam dryer Cyclones, heat 

exchange, tar 

reformer, & scrubber 

Stone & Webster, et al, (1995) Flue gas dryer Cyclones, heat 

exchange, & tar 

reformer 

Wan and Malcolm (1990) Flue gas dryer Cyclones, heat 

exchange & scrubber 

Weyerhaeuser (1992) Flue gas dryer Cyclones, heat 

exchange & scrubber 

Wright and Feinberg (1993) Unclear Quench system – 

details are not clear 

Heat exchange & 

solids – removal – 

Tar reformer system 

– details are not clear 

details are not 

clear 

Craig (1994) Rotary drum dryer Cyclones, heat 

exchange, & tar 

reformer 

20 
8.2 Other Capital Costs 

The cost of reactors, heat exchangers, compressors, blowers and pumps were determined 

using the energy and material balance from the Aspen Plus simulation along with the 

costing tool Questimate. The following were the sizing criteria. 

The reactors (ZnO, HTS, and LTS) were sized based on a gas hourly space velocity 

(GHSV), where GHSV is measured at standard temperature and pressure, 60°F and 1 atm 

(Fogler, 1992), and a height to diameter ratio of 2. The GHSV for the HTS and LTS 

reactor were set at 3,000/hr and 4,000/hr, respectively (typical values given in Kohl and 

Nielsen, 1997). The GHSV for each ZnO bed was set at 4,000/hr. 

The surface area of each heat exchanger was calculated based on the equation Q = 

U*A*∆Tln (where Q is the heat duty, U is the heat transfer coefficient, A is the 

exchanger surface area, and ∆Tln is the log mean temperature difference). Q was taken 

from the Aspen Plus simulation, U was estimated from literature sources (primarily 

Perry, et al, 1997), and ∆Tln was calculated using the temperatures in the Aspen Plus 

simulation. 

The design information including flow rate, operating temperature and pressure for the 

blowers and compressors were all taken from the Aspen Plus simulation. The cost of the 

syngas compressor (K-301) includes the cost of the interstage coolers and interstage 

knock out (K.O.) vessels. However, the cost of the interstage coolers for the hydrogen 

compressor (K-501) were not included in the Questimate cost estimate. Thus, these items 

had to be priced out separately. 

For the various pieces of equipment, the design temperature is determined to be the 

operating temperature plus 50°F (Walas, 1988). The design pressure is the higher of the 

operating pressure plus 25 psi or the operating pressure times 1.1 (Walas, 1988). 

The cost of the steam reformer was based on design and cost data in Leiby (1994). The 

reformer capital cost was determined and scaled based on heat duty. Literature values 

were also used to determine the capital and operating cost of the PSA unit (Schendel, et 

al, 1983 and Leiby 1994). The cost of the PSA unit was determined based on the 

hydrogen production rate. 

Some of the miscellaneous and balance of plant costs were scaled from information and 

costs in Aden, et al, (2002): 

- cooling tower 

- plant and instrument air 

- steam turbine/generator/condenser package 

- deaerator 

Appendix G: Equipment Design Parameters and Cost References contains the design 

parameters and cost references for the various pieces of equipment in the plant. 

21
9.0 Operating Costs 

There are two kinds of operating costs: variable and fixed costs. The following sections 

discuss the operating costs for the biomass gasification to hydrogen production plant 

including the assumptions and values for these costs. 

9.1 Variable Operating Costs 

There are many variable operating costs accounted for in this analysis. The variables, 

information about them, and costs associated with each variable are shown in Table 13. 

Table 13: Variable Operating Costs 

Variable Information and Operating Cost 

Tar reformer catalyst To determine the amount of catalyst inventory, the tar reformer was 

sized for a gas hourly space velocity (GHSV) of 2,000/hr based on the 

operation of the tar reformer at NREL’s TCPDU where GHSV is 

measured at standard temperature and pressure (Fogler, 1992). Initial 

fill then a replacement of 1% per day of the total catalyst volume. 

Price: $4.67/lb (Leiby, 1994) 

ZnO, steam reforming Initial fill then replaced every 5 years based on typical catalyst 

and shift catalyst lifetime. 

ZnO catalyst inventory based on GHSV of 4,000/hr. 

Steam reformer catalyst inventory based on inventory in Leiby, 1994 

and the ratio of the heat duty. 

Shift catalyst inventory based on GHSV of 3,000/hr for HTS and 

4,000/hr for LTS (typical values given in Kohl and Nielsen, 1997). 

Price (all three types): $4.67/lb (Leiby, 1994) 

Gasifier bed material Synthetic olivine and MgO. Delivered to site by truck equipped with 

self-contained pneumatic unloading equipment. Disposal by landfill. 

Olivine price: $172.90/ton (Jaekel, 2004) 

MgO price: $365/ton (Chemical Marketing Reporter, 2004) 

Solids disposal cost Price: $18/ton (Chem Systems Report, 1994) 

Electricity Price: 4.74¢/kWh (SRI, 2003) 

Natural gas Available at required pressure or pressure can be reduced. 

Temperature: 60°F 

Pipeline composition (mol%, dry) (Spath and Mann, 2000): 

CO2: 0.5% N2: 1.1% CH4: 94.4% C2H6: 3.1% 

C3H8: 0.5% i-C4H10: 0.1% n-C4H10: 0.1% C5 

+

: 0.2% 

H2S: 0.0004% 

Price: $5.28/MMBtu (SRI, 2003) 

Diesel fuel Usage: 10 gallon/hr plant wide use 

Price: $1.00/gallon (EIA, 2003) 

Chemicals Boiler chemicals – Price: $1.4/lb (Aden et al, 2003) 

Cooling tower chemicals – Price: $1.00/lb (Aden et al, 2003) 

LO-CAT chemicals – Price: $150/tonne of sulfur produced 

(Graubard, 2004) 

Waste Water The waste water is sent off-site for treatment. 

Price: $2.07/100ft3

 (East Bay Municipal Utility District, 2004) 

22 
9.2 Fixed Operating Costs 

Previous biomass gasification studies have not looked at fixed operating costs (i.e. 

salaries, overhead, maintenance, etc) in detail, therefore little data was available. As a 

result, the fixed operating costs given in Aden, et al, 2002 were used as a starting point to 

develop fixed costs for the biomass gasification-to-hydrogen production plant. Though 

hydrogen and ethanol production involve different processes and unit operations, it is 

reasonable as a first step to assume similar labor requirements because both designs are 

large-scale biomass conversion processes. However, this may be an area that would 

benefit from further examination by an engineering and consulting firm. 

The fixed operating costs used in this analysis are shown in Table 14 (labor costs) and 

Table 15 (other fixed costs). They are shown in 2002 U.S. dollars. The following 

changes in base salaries and number of employees were made compared to those used in 

the ethanol plant design in Aden, et al, 2002. 

•	 Plant manager salary raised from $80,000 to $110,000 

•	 Shift supervisor salary raised from $37,000 to $45,000 

•	 Lab technician salary raised from $25,000 to $35,000 

•	 Maintenance technician salary raised from $28,000 to $40,000 

•	 Shift operators salaries raised from $25,000 to $40,000 

•	 Yard employees salaries raised from $20,000 to $25,000 and number reduced 

from 32 to 12. 

•	 General manager position eliminated 

•	 Clerks and secretaries salaries raised from $20,000 to $25,000 and number

reduced from 5 to 3. 

The number of yard employees was changed to reflect a different feedstock and feed 

handling system compared to Aden, et al, 2002. Handling baled stover obviously 

requires more hands-on processing when compared to a wood chip feedstock. Based on a 

4-shift system, 3 yard employees were estimated to be needed, mostly to run the front end 

loaders. The general manager position was eliminated because a plant manager would 

likely be sufficient for this type of facility. Biomass gasification plants are more likely to 

operated by larger companies instead of operating like the dry mill ethanol model of 

farmer co-ops. Finally, the number of clerks and secretaries was reduced from 5 to 3. 

The estimate of three comes from needing 1 to handle the trucks and scales entering and 

leaving the facility, 1 to handle accounting matters, and 1 to answer phones, do 

administrative work, etc. 

Table 14: Labor Costs 

Position Salary Number Total Cost 

Plant manager $110,000 1 $110,000 

Plant engineer $65,000 1 $65,000 

Maintenance supervisor $60,000 1 $60,000 

Lab manager $50,000 1 $50,000 

23 
Position Salary Number Total Cost 

Shift supervisor $45,000 5 $225,000 

Lab technician $35,000 2 $70,000 

Maintenance technician $40,000 8 $320,000 

Shift operators $40,000 20 $800,000 

Yard employees $25,000 12 $300,000 

Clerks & secretaries $25,000 3 $75,000 

Total salaries (2002 $) $2,0800,000 

Since the salaries listed above are not fully loaded (i.e. do not include benefits), a general 

overhead factor was used. This also covers general plant maintenance, plant security, 

janitorial services, communications, etc. The 2003 PEP yearbook (SRI, 2003) lists the 

national average loaded labor rate at $37.66/hr. Using the salaries in Table 14 above 

along with the 60% general overhead factor from Aden, et al, 2002 gave an average 

loaded labor rate of $30/hr. To more closely match the PEP yearbook average, the 

overhead factor was raised to 95%. The resulting average loaded labor rate was $36/hr. 

Factors for maintenance, insurance, and taxes were obtained from Peters and 

Timmerhaus (2003). 

Table 15: Other Fixed Costs 

Cost Item Factor Cost 

General overhead 95% of total salaries $1,976,000 

Maintenance 2% of total project investment $3,072,500 

Insurance & taxes 2% of total project investment $3,072,500 

The updated salaries in Table 14 above were examined against salaries from a free salary 

estimation tool (BTA, 2004), which uses Bureau of Labor Statistics data and several 

other sources. Because the biomass analysis does not reflect a specific site in the United 

States, National Average Salaries for 2003 were used. With such an extensive listing of 

job titles in the salary estimation tool, a general position such as “clerks and secretaries” 

could be reflected by multiple job titles. In these instances, care was taken to examine 

several of the possible job titles that were applicable. A list of the job positions at the 

biomass-to-hydrogen production plant and the corresponding job titles in the salary 

estimation tool (BTA, 2004) is shown in Table 16. Overall, the salaries used in the 

biomass-to-hydrogen production plant design are close to the U.S. national average 

values given in column 4. 

24
Table 16: Salary Comparison 

Job Title in 

Biomass Plant 

Corresponding Job 

Title in Salary 

Estimating Tool (BTA 

2004) 

Salary Range (17th 

to 67th percentile) 

Average 

Salary (U.S. 

national 

average) 

Salary used in 

Biomass Plant 

Design (see 

Table 14) 

Plant manager Plant manager 

(experience) 

$81,042-$220,409 $106,900 $110,000 

Plant engineer Plant engineer $36,213-$66,542 $58,324 $65,000 

Maintenance 

supervisor 

Maintenance crew 

supervisor 

$35,036-$53,099 $45,191 $60,000 

Supervisor maintenance $34,701-$56,097 $47,046 

Supervisor maintenance 

& custodians 

$23,087-$45,374 $39,924 

Lab manager Laboratory manager $38,697-$70,985 $51,487 $50,000 

Shift supervisor Supervisor production $32,008-$51,745 $43,395 $45,000 

Lab technician Laboratory technician $25,543-$41,005 $34,644 $35,000 

Maintenance 

technician 

Maintenance worker $27,967-$46,754 $39,595 $40,000 

Shift operators Operator control room $33,983-$61,362 $49,243 $40,000 

Yard employees Operator front end loader $24,805-$39,368 $31,123 $25,000 

Clerks & Administrative clerk $19,876-$25,610 $26,157 $25,000 

secretaries Secretary $20,643-$31,454 $26,534 

Clerk general $15,984-$25,610 $22,768 

Overall, Aden, et al, 2002 lists fixed operating costs totaling $7.54MM in $2000. Using 

the labor indices, this equates to $7.85MM in $2002. On the other hand, the hydrogen 

design report has fixed operating costs totaling $10.2MM in $2002, which is $2.35MM 

higher. 

10.0 Pinch Analysis 

A pinch analysis was performed to analyze the energy network of the biomass 

gasification to hydrogen production process. The pinch technology concept offers a 

systematic approach to optimum energy integration of the process. First temperature and 

enthalpy data were gathered for the “hot” process streams (i.e., those that must be 

cooled), “cold” process streams (i.e., those that must be heated), and utility streams such 

as steam, flue gas, and cooling water. The minimum approach temperature was set at 50 

°F. A temperature versus enthalpy graph known as a composite curve was plotted for the 

hot and cold process streams. These two curves are shifted so that they touch at the pinch 

point. From this shifted graph, a grand composite curve is constructed which plots the 

enthalpy differences between the hot and cold composite curves as a function of 

temperature. This curve is shown in Figure 5 for the current design. This figure was 

used to determine the heat exchanger network of the system (Figure 4). 

25
Temperature Interval, F 

"Current" Design Grand Composite Curves 

Process 

2000 Fired Heater 

Cooling Water 1800 Air Cooler 

1600 BFW Preheat 

Steam Superhea

1400 Steam Generati

Degasifier Prehe

1200 1000 

800 600 400 200 

0 

-500,000,000 -400,000,000 -300,000,000 -200,000,000 -100,000,000 0 100,000,000 200,000,000 300,000,

Heat Flow, Btu/hr 

Figure 5: Current Design Grand Composite Curve 

26 
11.0 Energy Balance 

Because energy integration is so important to the hydrogen production process, 

understanding how and where the energy is utilized and recovered is key. Detailed 

energy balances around the major process areas were derived using data from the Aspen 

Plus simulation. Comparing the process energy inputs and outputs enables the energy 

efficiency of the process to be quantified. Also, tracing energy transfer between process 

areas makes it possible to identify areas of potential improvement. 

The philosophy of defining the “energy potential” of a stream is somewhat different from 

what was done for the ethanol process design report (Aden, et al, 2002). For that analysis 

the definition of the energy potential was based upon the higher heating values (HHVs) 

of each component. This HHV basis is convenient when a process is primarily made up 

of aqueous streams in the liquid phase. Since liquid water at the standard temperature has 

a zero HHV, the contributions for any liquid water is very small, especially as compared 

to any other combustible material also present in the stream. However, the hydrogen 

production process differs significantly in that most of the process streams are in the gas 

phase. To remove the background contributions of the water, the energy potential is 

instead based upon the lower heating values (LHVs) of each component. 

The total energy potential for a stream has other contributions beyond that of the heating 

value. Other energy contributions are: 

•	 Sensible heat effect – the stream is at a temperature (and pressure) different from 

that of the standard conditions at which the heating values are defined. 

•	 Latent heat effect – one or more components in the stream are in a different phase 

from that at which their heating values are defined. 

•	 Non-ideal mixing effect – any heating or cooling due to blending dissimilar 

components in a mixture. 

The procedure for actually calculating the energy potential of a stream is also different 

from what was done for the ethanol process design report. When the ethanol process was 

analyzed the contributions for the HHVs, sensible heating effects, and the latent heat 

effects were directly computed and combined. The calculations of the sensible and latent 

heat effects were done in an approximate manner. For example, the sensible heat effect 

was estimated from the heat capacity at the stream’s temperature, pressure, and 

composition; it was assumed that this heat capacity remained constant over the 

temperature range between the stream’s temperature and the standard temperature. 

However, the larger the difference between the stream temperature and the standard 

temperature, the more likely this assumption is not accurate. Indeed, the hydrogen 

production process operates at such large temperatures that this would not be an accurate 

way to account for the sensible heat effect. 

The enthalpy values reported by Aspen Plus can actually be adjusted in a fairly simple 

manner to reflect either an HHV or LHV basis for the energy potential. The enthalpies 

calculated and reported by Aspen Plus are actually based upon a heat of formation for the 

energy potential of a stream. So, the reported enthalpies already include the sensible, 

27
latent, and non-ideal mixing effects. If certain constants in Aspen’s enthalpy expressions 

could be modified to be based on either the components’ HHVs or LHVs instead of the 

heats of formation then Aspen Plus would report the desired energy potential values. 

However, since the constants cannot be easily changed, the reported enthalpy values were 

instead adjusted as part of a spreadsheet calculation. The factors used to adjust the 

reported enthalpies were calculated from the difference between each component’s heat 

of combustion (LHV) and the reported pure component enthalpy at combustion 

conditions. 

The major process energy inputs and outlets are listed in Table 17, along with their 

energy flowrates. Each input and output is also ratioed to the dry biomass energy 

entering the system. The biomass is of course the primary energy input, however other 

energy inputs are required. Natural gas is used as trim for the steam methane reformer, 

which is primarily fueled by the PSA offgas. Some electricity must be purchased from 

the grid to ensure that all power requirements are met. Air is also required for both the 

steam methane reformer as well as the char combustor, however it remains a minor 

energy input. Some water is used to wet the ash leaving the gasification system, 

however, the majority of process water is used for boiler feed water makeup and cooling 

water makeup. A large negative energy flow value is associated with this because it 

enters the process as a liquid. 

The sum of these energy outlets shown in Table 17 represents greater than 97% of the 

energy entering the system. The difference (< 3%) is comprised of energy losses due to 

ambient heating effects and work (pump, compressor) efficiency losses. 

Table 17: Current Design Overall Energy Analysis (LHV basis) 

 Energy Flow Ratio to Feedstock 

(MMBTU/hr, LHV basis) Energy Flow 

Energy Inlets 

Wood Chip Feedstock (dry) 1480.7 1.000 

Feedstock Moisture -209.7 -0.142 

Natural Gas 34.6 0.023 

Air 2.4 0.002 

Olivine 0.0 0.000 

MgO 0.0 0.000 

Water -268.7 -0.182 

Tar Reforming Catalyst 0.0 0.000 

Purchased Electricity 34.9 0.024 

Other 0.0 0.0 

Total 1074 0.725 

Energy Outlets 

Hydrogen 737.8 0.498 

Cooling Tower Evaporation 26.5 0.018 

Flue Gas 57.4 0.039 

Sulfur 0.6 0.000 

Compressor Heat 119.0 0.080 

Heat from Air-cooled Exchanger 149.3 0.101 

28 
 Energy Flow 

(MMBTU/hr, LHV basis) 

Ratio to Feedstock 

Energy Flow 

Ash 16.0 0.011 

Wastewater -18.7 -0.013 

Other -41.9 -0.028 

Total 1046 0.706 

The only saleable product from this process is hydrogen, but other important energy 

outlets also exist. There are two sources of flue gas: the char combustor and the steam 

methane reformer. Together, they total about 4% of the energy in the dried biomass. 

Cooling tower evaporative losses, wastewater, and ash are also minor energy outlets. 

However, two of the larger energy outlets come from air-cooled interstage cooling of the 

compressors, and from the air-cooling of the shifted syngas. Together, these two heat 

losses represent 18% of the energy that is not recovered within the process. Some of this 

heat could potentially be recovered using different heat exchange equipment, however it 

would likely be more expensive on an overall process basis to do so. 

The overall energy balance for the current design is depicted graphically in Figure 6. The 

energy values are listed as percentages of the dry biomass fed to the process. The 50% 

moisture entering the process within the wood chips has a negative value because it 

enters as a liquid. The same is also true for the negative values associated with cooling 

tower and steam cycle makeup water inputs (i.e. a latent heat “penalty”). 

Not all energy flows are shown within the context of this diagram. For example, the 

energy flows around the tar reforming and scrubbing section don’t appear to balance only 

because various integrated small streams are not shown in Figure 6. Crude syngas 

(83.3%) enters the section while wastewater (-1.3%), scrubbed syngas (73.4%), and 

cooling tower heat (2.6%) all exits. Thus there is a difference of 8.6% which is the heat 

going to the steam cycle that gets redistributed throughout the process. This heat 

integration does not appear directly on the diagram. This is also true for many of the 

other process areas. The heat integration, though not shown here, is depicted in an earlier 

diagram (Figure 4). 

It is also important to note that the 49.8% value listed for the hydrogen product should 

not be taken as the process efficiency. Instead, the summary sheet in Appendix A shows 

the hydrogen efficiency to be 45.6%. Remember that all energy inputs including 

electricity and natural gas must be factored into the process efficiency calculation even 

though these inputs are small. 

For comparison, the energy balance was also calculated on a HHV basis. This is shown 

in Figure 7. Some of the water streams are slightly negative due to the sensible heat 

effect. 

29
Interstage Natural 

Flue Gas Air Water MgO Olivine Compressor Air Gas Flue Gas Sulfur 

2% 

0.1% -0.1% 0.0% 0.0% Cooling 0.1% 2.3% 1.6% 0.0% 

Dry Biomass 7.4% 

100%

Steam Methane Feed Handling and Dried Crude Reformed S

Gasification Tar Reforming and Scrubbed Syngas Compression Compressed Reforming & Sulfur Conditioning Biomass Syngas Scrubbing Syngas Syngas Syngas 

Removal 

=99% 83.3% 73.4% 73.1% 82.0% -14% 

23.3% Biomass 16% 

Moisture Flue Gas PSA Offgas 

Ash Wastewater IP Steam 

1.1% -1.3% 

3.4% 

Feed H&D Work 0.2% Gasification Work 0.8% 

LP Steam 

Electricity 

Steam Cycle Generated Electricity Syngas Compr Work 5.0% 

0.4% Steam Reformer Work 0.1% 

H2 Compression Work 0.9%Cooling Tower Work 0.3%

6.2% 

Makeup Water-7.9% 

Purchased 2.4% Electricity 

Gasification Heat 0.2%

Tar Reform & Scrub Heat 2.6%

Compression Heat 0.2% CW Cooling Tower Evaporation 1.8% Shift & Purification Heat 0.6% 

H2 Compression Heat 0.1% Makeup Water -10.2% Steam Cycle Heat 5.8% 

Figure 6: Current Design Process Energy Balance (LHV Basis) 

30 
Interstage Natural 

Flue Gas Air Water MgO Olivine Compressor Air Gas Flue Gas Sulfur 

15% 

0.7% 0.0% 0.0% 0.0% Cooling 0.5% 2.4% 5.7% 0.0% 

Dry Biomass 6.8% 

100% 

Feed Handling and Dried Gasification Crude Tar Reforming and Scrubbed Steam Methane Reformed 

Conditioning Biomass Syngas Scrubbing Syngas Syngas Compression Compressed Syngas Reforming & Sulfur Syngas 

Removal


79.0% 75.9% -0.2% 89.7% 96.5% 100% 

25.1% Biomass 17% 

Moisture Flue Gas PSA Offgas 

Ash Wastewater IP Steam 

1.1% 0.3% 

15.5% 

Feed H&D Work 0.2% Gasification Work 0.8% 

LP Steam 

Electricity 

Steam Cycle Generated Electricity Syngas Compr Work 4.7% 

5.5% Steam Reformer Work 0.1% 

H2 Compression Work 0.8%Cooling Tower Work 0.2%

5.8% 

Makeup Water-0.1% 

Purchased 2.2%Electricity 

Gasification Heat 0.2%

Tar Reform & Scrub Heat 2.4%

Compression Heat 0.2% CW Cooling Tower Evaporation 8.8% Shift & Purification Heat 0.5% 

H2 Compression Heat 0.1% Makeup Water -0.2% Steam Cycle Heat 5.4% 

Figure 7: Current Design Process Energy Balance (HHV Basis) 

31 
12.0 Design, Modeling, and Capital Cost Changes for Goal Design 

The performance goals for the catalytic tar destruction and heteroatom removal work are 

shown in Table 18. Because the methane conversion is much higher than that for the 

current design (see Table 2), the process design was changed to eliminate the steam 

methane reformer. See Figure 3 for the block flow diagram and Appendix D: Goal 

Design Process Flow Diagrams for the PFDs. The main difference in the capital costing 

included the deletion of the steam methane reformer cost and the addition of a catalyst 

regenerator system and some cyclones. The heat for the reactor/regenerator system is 

supplied by combusting the PSA offgas in the regenerator vessel along with natural gas in 

order to operate the system isothermally. A breakdown of the capital costs for the goal 

design can be found in Appendix I: Goal Design Summary of Individual Equipment 

Costs. The rolled up TPI results were given previously in Table 10. 

Table 18: Goal Design Performance of Tar Reformer 

Compound Percent Conversion 

to CO & H2 

Methane (CH4) 80% 

Ethane (C2H6) 99% 

Ethylene (C2H4) 90% 

Tars (C10+) 99.9% 

Benzene (C6H6) 99% 

Ammonia (NH3)* 90% 

* Converts to N2 and H2 

Table 19 shows the operating parameters and outlet gas composition of the tar reformer 

for the goal design. More methane and higher hydrocarbons are reformed producing 

more hydrogen and carbon monoxide. The carbon monoxide is shifted to hydrogen after 

the sulfur removal step. 

32
--- 

Table 19: Goal Design Tar Reformer Properties and Outlet Gas Composition 

Tar reformer Variable Value 

Tar reformer inlet temperature 1,598oF (870oC) 

Tar reformer outlet temperature 1,598oF (870oC) 

Tar reformer outlet gas composition mol% (wet) mol% (dry) 

H2 41.62 53.18 

CO2 10.40 13.29 

CO 24.58 31.40 

H2O 21.73 

CH4 1.35 1.73 

C2H2 0.02 .02 

C2H4 0.19 0.24 

C2H6 0.001 0.002 

C6H6 0.0006 0.0007 

tar (C10H8) 0.0001 0.0001 

NH3 0.01 0.02 

H2S 0.03 0.04 

N2 0.06 0.08 

Gas heating value (Btu/lb) Wet: 5,311 HHV 4,794 LHV 

Dry: 6,960 HHV 6,282 LHV 

H2:CO molar ratio 1.69 

A breakdown of the power requirement for the goal design is given in Table 20. Again, 

this process design produces power but not enough to supply the electricity requirement 

of the plant. 

Table 20: Goal Design Plant Power Requirement 

Plant Section Power Requirement (kW) 

Feed handling & drying 742 

Gasification, Tar reforming/regeneration, 3,636 

& quench 

Compression & sulfur removal 26,058 

Shift, and PSA 159 

Hydrogen compression 4,190 

Steam system & power generation 662 required 

29,974 generated 

Cooling water & other utilities 1,152 

Miscellaneous 3,660 

Total plant power requirement 40,259 

Grid electricity requirement 10,284 

The heat integration of the system was reconfigured from the current design case. The 

resulting heat exchange network and pinch analysis for the goal design can be seen in 

Figure 8 and Figure 9, respectively. Additionally, the goal design energy balance on a 

LHV basis can be seen in Figure 10.

33 
Figure 8: Goal Design Heat Exchange Network within the Steam Cycle 

34 
Temperature Interval, F 

"Goal" Design Grand Composite Curves 

2000 

Process 

1800 Fired Heater 

Cooling Water 

1600 Air Cooler 

BFW Preheat 

1400 Steam Superhea

Steam Generatio

1200 

Degasifier Prehe

1000 

800 600 400 200 

0 

-500,000,000 -400,000,000 -300,000,000 -200,000,000 -100,000,000 0 100,000,000 200,000,000 300,000,

Heat Flow, Btu/hr 

Figure 9: Goal Design Grand Composite Curve 

35 
Natural Interstage 

Flue Gas Air Water MgO Olivine Gas Air Flue Gas Compressor Sulfur 

2% 

0.1% -0.1% 0.0% 0.0% 5.2% 0.1% 1.6% Cooling 0.0% 

Dry Biomass 8.7% 

100% Tar 

Feed Handling and Dried Crude Reforming/Catalyst Scrubbed Syngas Compression Compressed Reformed S

Gasification Sulfur Removal Conditioning Biomass Syngas Regneration and Syngas Syngas Syngas 

Scrubbing 99% 83.3% 78.7% 78.3% 81.6% -14% 

16%

Biomass 

Moisture Flue Gas Ash Wastewater

1.1% -0.5%

PSA Offgas 

IP S

20.7%

2.6% 

Feed H&D Work 0.2% 

Electricity 

Gasification Work 0.8% 

Generated 

Electricity LP Steam Syngas Compr Work 6.0% 

Steam Cycle 0.4% Shift & Purification Work 0.0% 

H2 Compression Work 1.0%Cooling Tower Work 0.3%

7.3% 

Makeup Water-7.1% 

Purchased 2.4% Electricity 

Gasification Heat 0.2%

Tar Reform & Scrub Heat 0.6%

Compression Heat 0.2% CW 

Shift & Purification Heat 0.5% Cooling Tower Evaporation 1.9% 

H2 Compression Heat 0.1% Makeup Water -10.6% Steam Cycle Heat 8.2% 

Figure 10: Goal Design Process Energy Balance (LHV Basis) 

36 
13.0 Resulting Economics of Current Design 

Once the capital and operating costs have been determined, a minimum hydrogen selling 

price (MHSP) can be determined using a discounted cash flow rate of return analysis. 

The methodology used is identical to that used in Aden, et al, (2002). The MHSP is the 

selling price of hydrogen that makes the net present value of the biomass syngas to 

hydrogen process equal to zero with a 10% discounted cash flow rate of return over a 20 

year plant life. An Excel worksheet was set up and some of the base case economic 

parameters used in the spreadsheet are given in Table 21. A sensitivity analysis was 

performed to examine the minimum hydrogen selling price for different debt/equity ratios 

at different internal rates of return (section 14.0 Current Design Sensitivity Analyses and 

section 16.0 Goal Design Sensitivity Analyses). 

Table 21: Economic Parameters 

Assumption Value 

Internal rate of return (after-tax) 10% 

Debt/equity 0%/100% 

Plant life 20 years 

General plant depreciation 200% DDB 

General plant recovery period 7 years 

Steam plant depreciation 150% DDB 

Steam plant recovery period 20 years 

Construction period 

1st 6 monts expenditures 

2.5 years 

8% 

Next 12 months expenditures 60% 

Last 12 months expenditures 32% 

Start-up time 6 months 

Revenues 50% 

Variable costs 75% 

Fixed costs 100% 

Working capital 5% of Total Capital Investment 

Land 6% of Total Purchased Equipment Cost 

(Cost taken as an expense in the 1st 

construction year) 

Note: The depreciation amount was determined using the same method as that 

documented in Aden, et al, 2002 using the IRS Modified Accelerated Cost Recovery 

System (MACRS). 

The resulting minimum hydrogen selling price for the current design is $1.38/kg 

($11.48/GJ, LHV) for a 2,000 bone dry tonne/day plant. A summary sheet of the capital 

and operating costs for the base case can be found in Appendix A: Current and Goal 

Base Case Summary Sheets. 

Figure 11 illustrates the cost contribution to product price for feedstock, capital, and 

operating costs by process area for this biomass gasification to hydrogen production 

process. Both percentages and contribution in terms of $/kg of hydrogen are given. The 

37
feedstock cost contributes the most to the product hydrogen price (31%). This is 

followed by gasification, tar reforming, and quench at 20%, compression and sulfur 

removal also at 20%, and steam reforming, shift, and hydrogen purification at 18%. 

Although the system produces power, it does not produce enough to meet the plant’s 

internal power requirements. The steam cycle generates almost 26 MW of power but the 

plant requires almost 36 MW of power, largely due to the syngas compression 

requirement. Thus 10 MW of power is purchased from the grid. 

14.0 Current Design Sensitivity Analyses 

Many sensitivity cases were run to examine the effects of several parameters on the 

current base case design Table 22 outlines the different sensitivity cases that were 

examined. Table 23 contains the results for the sensitivity analysis and Figure 12 shows 

the results in Table 23 graphically. Internal rate of return (IRR) and debt equity ratio 

were also examined. When a percentage of the financing is debt, the loan interest rate 

was set at 7.5% with a loan term of 10 years. Figure 13 is a graph showing those results 

and how the minimum hydrogen selling price changes with different combinations of 

IRR and debt/equity. 

38
ing

ifi

i

l

ion 

ies 

ials l ici

ici Fi

$0.13 

$0.04 

(10%) 

(20%) 

) 

) 

) 

) 

) 

Feedstock 

Feed Handling and Dry

 Gas cation, Tar 

Reforming, & Quench 

Compress on and Sulfur 

Remova

 Steam Methane 

Reforming, Shift, and PSA

 Hydrogen Compress

Steam System and 

Power Generation

 Cooling Water and Other 

Utilit

Capital Recovery Charge Raw Mater Process E ectr ty 

Electr ty Generated xed Costs 

$0.43 

$0.14 

$0.27 

$0.27 

$0.25 

$0.05 

-$0.19 (-4%) Net 

MHSP = 

$1.38/kg (3%

(20%

(18%

(3%

(31%

-$0.30 -$0.20 -$0.10 $0.00 $0.10 $0.20 $0.30 $0.40 $0.5

Figure 11: Current Design Base Case Cost Contribution Diagram 

39
Table 22: Current Design - Sensitivity Analysis Cases 

Letter Sensitivity Case Analysis Changes Made 

A Decrease feedstock cost 

to $0/dry ton 

The feedstock cost in the DCFROR spreadsheet was changed from 

$30/dry ton to $0/dry ton. 

B Increase feedstock cost 

to $53/dry ton 

The feedstock cost in the DCFROR spreadsheet was changed from 

$30/dry ton to $53/dry ton. 

C Lower feed moisture 

content of 30 wt% 

The feed moisture content in the Aspen Plus model was decreased 

from 50 wt% to 30 wt%. 

D Less drying of biomass 

feed to a moisture 

content of 20 wt% 

The wood moisture content at the dryer outlet was changed from 12 

wt% to 20 wt%. The gasifier temperature dropped from 870oC 

(1,598oF) to 859oC (1,578oF). No additional natural gas was required 

to maintain the heat balance around the gasifier and combustor 

(enough additional char was produced at the lower gasifier 

temperature). The dryer cost decreased. 

E Less drying of biomass 

feed to a moisture 

content of 20 wt% and 

keep the gasifier 

temperature constant 

The wood moisture content at the dryer outlet was changed from 12 

wt% to 20 wt%. The olivine circulating between the gasifier and 

combustor had to be increased by a factor of 1.12 times the base case 

flow to maintain a gasifer temperature of 870oC (1,598oF). Natural 

gas at a rate of 1,709 lb/hr was added to the combustor in order to 

maintain the heat balance around the gasifier and combustor. The 

dryer cost decreased and the gasifier/combustor cost increased. 

F Less drying of biomass 

feed to a moisture 

content of 30 wt% 

The wood moisture content at the dryer outlet was changed from 12 

wt% to 30 wt%. The gasifier temperature dropped from 859oC 

(1,598oF) to 870oC (1,547oF). Natural gas at a rate of 3,417 lb/hr was 

added to the combustor in order to maintain the heat balance around 

the gasifier and combustor. The dryer cost decreased. 

G Less drying of biomass 

feed to a moisture 

content of 30 wt% and 

keep the gasifier 

temperature constant 

The wood moisture content at the dryer outlet was changed from 12 

wt% to 30 wt%. Olivine circulating between the gasifier and 

combustor increased by a factor of 1.3 times the base case flow to 

maintain a gasifer temperature of 870oC (1,598oF). Natural gas at a 

rate of 8,543 lb/hr was added to the combustor in order to maintain 

the heat balance around the gasifier and combustor. The dryer cost 

decreased and the gasifier/combustor cost increased. 

H No dryer The dryer was removed from the Aspen Plus model. The olivine 

circulating between the gasifier and combustor had to be increased by 

a factor of 1.9 times the base case flow to maintain a gasifer 

temperature of 870oC (1,598oF). Natural gas at a rate of 23,920 lb/hr 

was added to the combustor in order to maintain the heat balance 

around the gasifier and combustor. The dryer cost was eliminated. 

The gasifier/combustor cost increased. There is a net power 

generation of 34 MW from the system instead of a deficiency of 10 

MW which had to be purchased from the grid for the base case. 

I Lower gasifier 

steam:wood ratio of 0.1 

and keep the gasifier 

temperature constant 

The steam:wood ratio to the gasifier was decreased from 0.4 to 0.1. This 

lower rate was based on the operation of the gasifier at Burlington, 

Vermont during sustained operation and testing for this demonstration 

project (Overend, 2004). The olivine circulating between the gasifier and 

combustor was decreased by a factor of 0.87 times the base case rate to 

maintain a gasifer temperature of 870oC (1,598oF). The 

gasifier/combustor cost decreased. 

J Higher gasifier 

steam:wood ratio of 1 

The steam:wood ratio to the gasifier was increased from 0.4 to 1. The 

olivine circulation rate was kept the same as the base case and thus the 

gasifier temperature decreased from 870oC (1,598oF) to 847oC (1,557oF). 

Natural gas at a rate of 1,709 lb/hr was added to the combustor in order to 

maintain an energy balance around the gasifier and combustor. The 

gasifier/combustor cost increased. 

40
Letter Sensitivity Case Analysis Changes Made 

K Higher gasifier 

steam:wood ratio of 1 

and keep the gasifier 

temperature constant 

The steam:wood ratio to the gasifier was increased from 0.4 to 1. 

Typically, direct gasifiers operate at a steam:wood ratio closer to 1. 

However, this rate was tested here to determine the effects on the 

indirect gasifier system. The olivine circulating between the gasifier 

and combustor had to be increased by a factor of 1.25 times the base 

case rate to maintain a gasifer temperature of 870oC (1,598oF). 

Natural gas at a rate of 5,467 lb/hr was added to the combustor in 

order to maintain an energy balance around the gasifier and 

combustor. The gasifier/combustor cost increased. 

L No H2 recycle to PSA The recycling of hydrogen to the PSA feed was eliminated. 

M Eliminate LTS The LTS was removed from the Aspen Plus model. The LTS cost 

was eliminated. 

N Lower tar reformer 

catalyst replacement 

The tar reformer catalyst replacement was lowered from 1 vol% to 0.5 

vol%. 

O Treat waste water 

internally 

Instead of sending the waste water stream off-site for treatment. A 

reverse osmosis system was installed at the plant. The waste water 

was cleaned and sent to the steam cycle. 

P Increase in PSA cost There is some variability in the capital cost data for the PSA so the 

cost was increased by a factor of 1.6 to determine the sensitivity to 

this parameter. This factor was determined using two different 

costing methods for the PSA. One was based on the hydrogen 

production rate and the other was based on the inlet flow rate to the 

PSA. 

Q Increase in steam 

reforming cost 

There is some variability in the capital cost data for the steam 

reformer so the cost was increased by a factor of 2 to determine the 

sensitivity to this parameter. The cost of the steam reformer was 

based on the duty but there could be some deviation from a standard 

steam methane reformer because the stream being reformed contains a 

low concentration of methane. 

R Increase in electricity 

price to 6¢/kWh 

The electricity price in the DCFROR spreadsheet was changed from 

4.74¢/kWh to 6¢/kWh. 

S Increase in natural gas 

price to $7/MMBtu 

The natural gas cost in the DCFROR spreadsheet was changed from 

$5.28/MMBtu to $7/MMBtu. 

T Lower feed handling & 

drying capital cost 

The feed handling and drying cost was reduced from the average cost 

in Table 11 to the second lowest cost in Table 11. 

U Lower gasification & 

clean up capital cost 

The gasification and gas clean up cost was reduced from the average 

cost in Table 11 to the second lowest cost in Table 11. 

V Combined lower feed 

handling & drying and 

lower gasification & 

clean up capital cost 

Both the feed handling and drying cost and the gasification and gas 

clean up cost were reduced to the second lowest cost in Table 11. 

W Higher feed handling & 

drying capital cost 

The feed handling and drying cost was increased from the average 

cost in Table 11 to the second highest cost in Table 11. 

X Higher gasification & 

clean up capital cost 

The gasification and gas clean up cost was increased from the average 

cost in Table 11 to the second highest cost in Table 11. 

Y Combined higher feed 

handling & drying and 

higher gasification & 

clean up capital cost 

Both the feed handling and drying cost and the gasification and gas 

clean up cost were increased to the second highest cost in Table 11. 

41
Table 23: Current Design - Base Case and Sensitivity Analysis Results 

Letter Sensitivity Case Minimum 

Hydrogen Selling 

Price ($/kg) 

Minimum 

Hydrogen Selling 

Price ($/GJ, LHV) 

Base Current design - base case $1.38 $11.48 

A Decrease feedstock cost to $0/dry ton $0.94 $7.86 

B Increase feedstock cost to $53/dry ton $1.71 $14.24 

C Lower feed moisture content of 30 wt% $1.31 $10.89 

D Less drying of biomass feed to a moisture content of 

20 wt% 

$1.37 $11.44 

E Less drying of biomass feed to a moisture content of 

20 wt% and keep the gasifier temperature constant 

$1.39 $11.59 

F Less drying of biomass feed to a moisture content of 

30 wt% 

$1.46 $12.20 

G Less drying of biomass feed to a moisture content of 

30 wt% and keep the gasifier temperature constant 

$1.50 $12.50 

H No dryer $1.78 $14.85 

I Lower gasifier steam:wood ratio of 0.1 and keep the 

gasifier temperature constant 

$1.30 $10.87 

J Higher gasifier steam:wood ratio of 1 $1.57 $13.07 

K Higher gasifier steam:wood ratio of 1 and keep the 

gasifier temperature constant 

$1.58 $13.19 

L No hydrogen recycle to PSA $1.30 $10.87 

M Eliminate LTS $1.47 $12.23 

N Lower tar reformer catalyst replacement of 0.5 vol% $1.35 $11.27 

O Treat waste water internally $1.38 $11.49 

P Increase in PSA cost $1.42 $11.82 

Q Increase in steam reforming cost $1.45 $12.07 

R Increase in electricity price to 6¢/kWh $1.40 $11.64 

S Increase in natural gas price to $7/MMBtu $1.39 $11.55 

T Lower feed handling & drying capital cost $1.35 $11.24 

U Lower gasification & clean up capital cost $1.35 $11.22 

V Combined lower feed handling & drying and lower 

gasification & clean up capital cost 

$1.32 $10.99 

W Higher feed handling & drying capital cost $1.41 $11.78 

X Higher gasification & clean up capital cost $1.42 $11.85 

Y Combined higher feed handling & drying and higher 

gasification & clean up capital cost 

$1.46 $12.15 

42
$0.80 

$1.00 $1.20 $1.40 $1.60 $1.80 $2.00

Minimum Hydrogen Selling Price ($/kg) 

Base "Current" 

$0 feed 

$53 feed 

30% moist feed 

dry to 20% moist 

dry to 20% moist; same T 

dry to 30% moist 

dry to 30% moist; same T 

dry to 40% moist 

dry to 40% moist; same T 

no dryer 

stm:wood = 0.1 

stm:wood = 1; same T 

stm:wood = 1 

no H2 recycle 

Internal WWT

no LTS 

0.5% cat replace 

Increase PSA cost 

Increase in steam reformer cost 

Higher electricity cost 

Higher nat gas cost 

Low feed & dry cost 

Low gas & clean co

Low cost feed thru c

High 

Hig

High fe

Figure 12: Current Design Sensitivity Analysis Results 

43
Figure 13: Effect of IRR and Debt/Equity on Current Design Base Case 

Since the feedstock cost contributes a large percentage to the hydrogen selling price, the 

zero feedstock case (A) results in the lowest hydrogen price. Conversely, increasing the 

feedstock cost (B) adversely affects the minimum hydrogen selling price. 

The no dryer case (H) results in the highest hydrogen selling price. In this case, 

eliminating the dryer eliminates the capital cost for that piece of equipment. 

Additionally, because there is excess high temperature heat available that would have 

been used for drying, this case results in more electricity being produced than consumed 

by the plant. However, the size and thus cost of the gasifier/combustor system increases 

and the amount of natural gas that must be added to the combustor is significant resulting 

in a hydrogen selling price that is higher than the base case. 

Significantly increasing or decreasing the gasifier steam to wood ratio (I, J, and K) has a 

large affect on the minimum hydrogen selling price. This variable greatly affects on the 

heat balance of the system and the capital and operating costs. 

Feeding a lower moisture feedstock (C) also affects the heat balance, thus resulting in a 

decrease in the hydrogen price. More heat is available for power production. 

Less drying of the biomass (i.e., a higher moisture content biomass exiting the dryer) was 

also examined in the sensitivity analysis. Instead of drying to 12%, the biomass was 

dried to a moisture content of 20% in two cases (D and E) and to a moisture content of 

30% in two other cases (F and G). Although less drying affects the heat balance of the 

system, drying to a moisture content of 20% (D and E) resulted in virtually the same 

hydrogen selling price as the base case. For the case where the gasifier temperature is 

kept constant (E), the hydrogen price does not decrease from the base case because there 

44
is an increase in operating costs (natural gas must be added to the combustor) even 

though the total project investment decreases slightly. For the case where the gasifier 

temperature is reduced (D), the hydrogen yield decreases and there is a decrease in the 

total project investment. However, drying to a moisture content of 30% did increase the 

minimum hydrogen selling price (F and G). This is due to decreased hydrogen yields and 

increased operating costs (natural gas) in both cases (F and G) and an increase in the total 

project investment for the constant gasifier temperature case (G). It should be noted that 

both of these cases (F and G) did generate more electricity than what was required for the 

plant. 

A general observation can be made about the differences between sensitivity case D and 

E, between sensitivity case F and G, and between sensitivity case J and K. In all three of 

these instances lowering the gasifier temperature decreases the hydrogen yield but adding 

natural gas to the combustor along with increasing the olivine circulating rate will 

increase the gasifier temperature. However, the increase in operating cost coupled with 

any capital cost increases for case E, G, and K is slightly more detrimental than the lower 

hydrogen yield for case D, F, and J. 

The case of eliminating the LTS reactor (M) was examined because often plants with 

PSA units will use only a HTS reactor followed by a PSA. This is because the PSA can 

easily remove CO and other components to produce a high purity hydrogen stream. 

Eliminating the LTS reactor (M) increases the hydrogen price because of a reduction in 

hydrogen yield that is not recovered by the increase in electricity produced. The LTS 

reactor is a low capital cost item. Although the PSA can easily remove CO and other 

components to produce a high purity hydrogen stream, in this case, it is more economical 

to leave the LTS reactor in. 

Assuming a hydrogen recovery rate of 85% without recycling a portion of the product 

hydrogen to the inlet of the PSA (L) results in a higher hydrogen yield and thus a lower 

minimum hydrogen selling price. Although increasing the PSA cost (P) did increase the 

hydrogen price it did not have as large of an effect as the no hydrogen recycling case. 

Increasing the steam reformer cost (Q) increased the minimum hydrogen selling price. 

This capital cost along with the PSA capital cost are items where vendor quotes would 

reduce the uncertainty in these larger capital cost items. 

Because the feed handling and drying costs as well as the gasification and gas clean up 

costs came from cost data in other detailed studies there is a larger amount of uncertainty 

as to the exact costs that should be used in this process design. Therefore, several 

sensitivity cases were run for lower and higher capital costs for the feed handling and 

drying section and for the gasification and gas clean up section. Overall, decreasing the 

costs to the second lowest cost from the various studies (T and U) reduced the minimum 

hydrogen selling price but not significantly, only about 2%. Additionally, increasing the 

costs to the second highest cost (W and X) did not increase the hydrogen price 

considerably, only about 3%. A combination of increasing and decreasing the capital 

cost for both the feed handling and drying section and the gasification and gas clean up 

45
section was also tested (Y). This had a larger effect on the change in the minimum 

hydrogen selling price. The price decreased from $1.38/kg to $1.32/kg for the low 

capital cost case (V) and the price increased from $1.38/kg to $1.46/kg for the high 

capital cost case (Y). 

Treating the waste water stream internally (O) had virtually no effect on the overall 

economics. Three cases that had very little effect on the minimum hydrogen selling price 

are decreasing the amount of tar reformer catalyst that must be replaced (N), increasing 

the electricity price (R), and increasing the natural gas price (S). This is because all of 

these items contribute a small amount to the overall operating cost. 

15.0 Resulting Economics of Goal Design 

The resulting minimum hydrogen selling price for the goal design is $1.24/kg 

($10.34/GJ, LHV) for a 2,000 bone dry tonne/day plant. The hydrogen price decreases 

from the current base case design (which is $1.38/kg or $11.48/GJ, LHV) mainly because 

of an increase in the hydrogen yield. The decrease in the total project investment has 

some effect. A summary sheet of the capital and operating costs for the base case can be 

found in Appendix A: Current and Goal Base Case Summary Sheets. The cost 

contribution to product price for feedstock, capital, and operating costs by process area 

for the goal design can be seen in Figure 14. Both percentages and contribution in terms 

of $/kg of hydrogen are given. Again, the feedstock cost contributes the most to the 

product hydrogen price (32%) and although the system produces power, it does not 

produce enough to meet the plant’s internal power requirements. Comparing the cost 

contribution of the goal design (Figure 14) with that for the current design (Figure 11) 

shows an increase in the gasification/tar reforming/regeneration/quench bar and a 

decrease in the shift/PSA bar. This happens because the capital and operating costs 

associated with the steam methane reformer are removed. However, there are capital and 

operating costs associated with adding the tar catalyst regenerator. 

46
li

i

i / 

l

l

i

ion

i

i

l i

El Fi

) 

) 

) 

Feedstock 

Feed Hand ng and 

Drying

 Gas fication, Tar 

Reform ng

Regeneration, & Quench 

Compression and Sufur 

Remova

 Sh ft, and PSA

 Hydrogen Compress

 Steam System and 

Power Generat on

 Cool ng Water and 

Other Utilities 

Capita Recovery Charge Raw Mater als Process Electricity 

ectricity Generated xed Costs 

$0.13 

$0

$0.28 

$0.11 

$0.05 

$0.12 

$0.03 

-$0.20 

(10%

(-6%) Net 

(9%

(4%) 

(3%

-$0.30 -$0.20 -$0.10 $0.00 $0.10 $0.20 $0.30 

Figure 14: Goal Design Base Case Cost Contribution Diagram 

47 
16.0 Goal Design Sensitivity Analyses 

Only a few of the parameters tested in the current design sensitivity analysis were tested 

here. Overall, the parameters tested on the current design will have a similar affect on the 

goal design. Since the feedstock cost has a big impact on the hydrogen price, the lower 

and higher feedstock costs were tested here. Because the natural gas consumption 

increased compared to the current design, the effect of increasing the cost of natural gas 

was also examined. A few of the other parameters listed above in the current design 

sensitivity analysis (Table 22) were also tested and are shown in Table 24. Additionally, 

changing the steam to carbon ratio to the shift reactors was investigated. All of the 

variables examined in the sensitivity analysis for the goal design are listed in Table 24 

and the results are in Table 25. The results of the sensitivity analysis for the goal design 

are also shown in Figure 15. Internal rate of return and debt equity ratio were also 

examined for the goal design. Again, when a percentage of the financing is debt, the loan 

interest rate was set at 7.5% with a loan term of 10 years. Figure 16 shows those results. 

Table 24: Goal Design – Sensitivity Analysis Cases 

Letter Sensitivity Case Analysis Changes Made 

AA Decrease feedstock cost to 

$0/dry ton 

The feedstock cost in the DCFROR spreadsheet was changed from 

$30/dry ton to $0/dry ton. 

BB Increase feedstock cost to 

$53/dry ton 

The feedstock cost in the DCFROR spreadsheet was changed from 

$30/dry ton to $53/dry ton. 

CC Lower feed moisture content 

of 30 wt% 

The feed moisture content in the Aspen Plus model was decreased 

from 50 wt% to 30 wt%. 

DD Less drying of biomass feed 

to a moisture content of 20 

wt% 

The wood moisture content at the dryer outlet was changed from 12 

wt% to 20 wt%. The gasifier temperature dropped from 859oC 

(1,598oF) to 870oC (1,578oF). No additional natural gas was 

required to maintain the heat balance around the gasifier and 

combustor (enough additional char was produced at the lower 

gasifier temperature). The dryer cost decreased. 

EE No hydrogen recycle to PSA The recycling of hydrogen to the PSA feed was eliminated. 

FF Increase in PSA cost There is some variability in the capital cost data for the PSA so the 

cost was increased by a factor of 1.6 to determine the sensitivity to 

this parameter. This factor was determined using two different 

costing methods for the PSA. One was based on the hydrogen 

production rate and the other was based on the inlet flow rate to the 

PSA. 

GG Increase in natural gas price 

to $7/MMBtu 

The natural gas cost in the DCFROR spreadsheet was changed from 

$5.28/MMBtu to $7/MMBtu. 

HH Increase in tar 

reformer/catalyst regenerator 

system capital cost 

The capital cost for the tar reformer/regenerator system was doubled 

making the total project investment of the goal base case design 

roughly the same as that for the current base case design. 

II Increase in shift steam to 

carbon ratio from 2 to 3 

The shift steam rate in the Aspen Plus was increased from a 

steam:carbon ratio of 2 mol H2O/mol of C to a value of 3. 

JJ Decrease in shift steam to 

carbon ratio from 2 to 1.5 

The shift steam rate in the Aspen Plus was decreased from a 

steam:carbon ratio of 2 mol H2O/mol of C to a value of 1.5. 

48 
Figure 15: Goal Design Sensitivity Analysis Results 

Figure 16: Effect of IRR and Debt/Equity on Goal Design Base Case 

49
Table 25: Goal Design Base Case and Sensitivity Analysis Results 

Letter Sensitivity Case Minimum 

Hydrogen Selling 

Price ($/kg) 

Minimum 

Hydrogen Selling 

Price ($/GJ, LHV) 

Base Goal design - base case $1.24 $10.34 

AA Decrease feedstock cost to $0/dry ton $0.84 $6.97 

BB Increase feedstock cost to $53/dry ton $1.55 $12.9 

CC Lower feed moisture content of 30% $1.18 $9.81 

DD Less drying of biomass feed to a moisture content of 

20 wt% 

$1.26 $10.47 

EE No hydrogen recycle to PSA $1.21 $10.08 

FF Increase in PSA cost $1.28 $10.67 

GG Increase in natural gas price to $7/MMBtu $1.26 $10.49 

HH Increase in tar reformer/catalyst regenerator system 

capital cost 

$1.27 $10.6 

II Increase in shift steam to carbon ratio from 2 to 3 $1.28 $10.63 

JJ Decrease in shift steam to carbon ratio from 2 to 1.5 $1.22 $10.21 

Even increasing the capital cost of the tar reformer/regenerator system so that the total 

project investment was equivalent to that of the current design (HH) resulted in a 

minimum hydrogen selling price that is less than the minimum hydrogen selling price for 

the current base case design. This is because the hydrogen yield for this design is higher. 

A higher steam to carbon ratio increases the hydrogen yield but adversely affects the 

economics of the goal design because the operating costs increase and the total project 

investment goes up as well. However, there is a minimum steam to carbon ratio that the 

system must operate at in order to convert the CO to hydrogen (CO + H2O Ù CO2 + H2). 

17.0 Sensitivity to Plant Size 

The plant size is another variable that was examined for both the current and goal case 

design. The plant size was changed in the spreadsheet from the base case size of 2,000 

dry tonne/day to the desired plant size. The material and energy balances were 

determined by multiplying the base case values by the ratio of the plant sizes (i.e., 

multiplying by [the desired plant size in dry tonne/day]/[2,000 dry tonne/day]). The 

equipment were then scaled using the scaling exponents shown in Appendix H: Current 

Design Summary of Individual Equipment Costs and Appendix I: Goal Design Summary 

of Individual Equipment Costs (i.e., new cost = original cost * [new size/original size]exp) 

and the minimum hydrogen selling price was recalculated. Figure 17 shows the 

difference in the minimum hydrogen selling price for a plant size of 500 bone dry 

tonnes/day to 2,000 bone dry tonnes/day. In reducing the plant size from 2,000 bone dry 

tonnes/day to 500, the hydrogen price increases from $1.38/kg to $1.88/kg for the current 

design and from $1.24/kg to $1.68/kg for the goal design. This is a 36% increase. 

50
Figure 17: Effect of Plant Size on Minimum Hydrogen Selling Price 

18.0 Syngas Price 

As can be seen in Figure 11 and Figure 14, syngas production accounts for a significant 

portion of the minimum hydrogen selling price. This is also true for the synthesis of 

other fuel or chemical products (Spath and Dayton, 2003). As a benchmark for 

thermochemical conversion, the DOE Biomass Program is setting program targets based 

on intermediate syngas prices to track progress toward reducing the technical barriers 

associated with biomass gasification. Therefore, this analysis included calculations in 

determining both an intermediate and a stand-alone clean, reformed syngas price. 

18.1 Intermediate Syngas Price 

First an intermediate syngas price was determined. The value of the syngas was 

determined by taking a slipstream of the clean, reformed syngas and treating it as a minor 

co-product to the overall biomass-to-hydrogen process. The price of the syngas 

slipstream was determined to be the value that would maintain the MHSP equal to that of 

the base case hydrogen price which does not have a slipstream. This was done by taking 

the Aspen Plus model and separating a slipstream of clean, reformed syngas from the 

process, setting the hydrogen price equal to the base case cost (i.e., $1.38/kg for the 

current design and $1.24/kg for the goal design), and calculating the syngas price using 

the revised material and energy balance and thus revised capital and operating costs. 

In order to calculate an intermediate syngas price, a slipstream of clean, reformed syngas 

from 1%-20% of the total syngas stream was examined. The heat balance was the 

limiting factor beyond 20%, resulting in no flow through the steam cycle beyond the 

steam required for gasification and reforming. A slipstream larger than this amount 

would require the combustion of natural gas or another fuel to raise steam. The 

slipstream for the current design was taken just downstream of the steam reformer (R51
401). Since the goal design eliminates the steam reformer, the slipstream for the goal 

design was taken just after the ZnO beds (R-302). Therefore, both of these systems are 

examining clean, reformed syngas. The intermediate syngas price in $/GJ (LHV) for 

both designs can be seen in Figure 18. For the current design the intermediate syngas 

price starts out at $6.88/GJ ($7.25/MMBtu) for a 1% slipstream and ramps up to 

$8.24/GJ ($8.69/MMBtu) for a 20% slipstream. In the goal design the intermediate 

syngas price starts out at $4.98/GJ ($5.25/MMBtu) for a 1% slipstream and ramps up to 

$6.97/GJ ($7.35/MMBtu) for a 20% slipstream. The intermediate syngas price of the 

clean, reformed syngas for the integrated process should actually be considered to be the 

low end value at the small slipstream amount. This is the cost of the syngas for the 

integrated process. As the slipstream becomes larger, the price escalates quickly and then 

levels off thus approaching the syngas price of a stand-alone plant (see section 18.2 

Stand-alone Syngas Price). 

Figure 18: Intermediate Syngas Price 

18.2 Stand-alone Syngas Price 

Next a stand-alone syngas price was determined. For the current case this meant 

removing the process steps downstream of the steam reformer (shift conversion, 

purification, and hydrogen compression), and reconfiguring the heat balance. For the 

goal case this meant removing the process steps downstream of the sulfur removal step 

(shift conversion, purification, and hydrogen compression) and reconfiguring the heat 

balance. The syngas is cooled and the water is condensed from the syngas stream but no 

other conditioning of the syngas is done. 

In the current and goal case integrated hydrogen production process designs, off gas from 

the PSA unit is used to fuel the steam reformer or tar regenerator, respectively, with a 

slight amount of natural gas used for combustion control. In the stand-alone syngas plant 

for the current and goal designs, only natural gas is used as fuel since the product is now 

52
syngas. The heat available within the stand-alone syngas plant is used to meet the steam 

demand of the system, which means steam required for gasification and for the current 

design, additional steam required for steam methane reforming. Some power is also 

produced. The resulting stand-alone syngas price for each design is given in Table 26.

Table 26: Stand-alone Syngas Price 

Current Design Goal Design 

Stand-alone syngas price (LHV) $8.22/GJ $6.73/GJ 

$8.67/MMBtu $7.10/MMBtu 

For each stand-alone syngas design compared to the integrated hydrogen production 

plant, the total project investment decreases but the operating cost for natural gas and 

electricity increases. However, the natural gas and electricity operating costs for the 

stand-alone syngas goal design do not increase as much as those for the stand-alone 

syngas current design. This is because the shift conversion section has been eliminated 

and thus for the stand-alone goal design there is no additional steam requirement other 

than that for gasification. 

19.0 Hydrogen Program Analysis 

The results of this analysis are being used by the US Department of Energy’s Hydrogen, 

Fuel Cells & Infrastructure Program in the standard worksheet that they have developed 

for their hydrogen analysis group. However, it should be noted that the hydrogen price 

determined from their spreadsheet will be different than ours due to their use of different 

economic parameters such as operating hours, feedstock cost, inflation and escalation. It 

should also be noted that the Hydrogen, Fuel Cells & Infrastructure Program funded a 

portion of this work. 

20.0 Conclusions 

The results of this analysis show a minimum hydrogen selling price of $1.38/kg 

($11.48/GJ, LHV) for a 2,000 bone dry tonne/day plant for the current design and a price 

of $1.24/kg ($10.34/GJ, LHV) for the goal design. The hydrogen price decreases mainly 

because of an increase in the hydrogen yield. The decrease in the total project investment 

also has some affect. This result shows that the research at NREL in catalytic tar 

destruction and heteroatom removal is moving in a direction that has the potential to 

decrease the cost of producing clean syngas (by about $1.5-2/GJ) and any subsequent fuel 

products via biomass gasification. 

Since the feedstock cost contributes a large percentage to the hydrogen selling price 

(about 30%), this variable will always have a large impact on the economics. Overall, the 

sensitivity analysis shows that any parameter that significantly affects the heat balance of 

the system will greatly affect the minimum hydrogen selling price. For example, 

eliminating the dryer and adding more natural gas to the char combustor eliminates the 

dryer capital cost but increases operating costs and capital costs associated with the 

gasifier/combustor in order to maintain the heat balance around the gasifier/combustor. 

53
Also, significantly increasing or decreasing the gasifier steam to wood ratio has a large 

affect on the minimum hydrogen selling price. This variable greatly affects on the heat 

balance of the system and the capital and operating costs. Feeding a lower moisture 

feedstock (the base case assumes 50% moisture in the feed) also affects the heat balance, 

thus resulting in a decrease in the hydrogen price. 

The intermediate syngas price for the current and goal designs are $6.88/GJ 

($7.25/MMBtu) and $4.98/GJ ($5.25/MMBtu), respectively. This is for clean, reformed 

syngas in the integrated biomass-to-hydrogen design. Stand-alone syngas plants are not 

being built today but for a stand-alone plant the syngas price would be $8.24/GJ 

($8.69/MMBtu) for a plant based on the current design and $6.97/GJ ($7.35/MMBtu) for 

a plant based on the goal design. The lower intermediate syngas price shows the 

importance of integration within the fuels synthesis process plant. 

21.0 Future Work 

In addition to gas clean up and conditioning other barrier areas that could reduce the cost 

of fuel products from thermochemical conversion of biomass include feed handling and 

drying, gasification, production of different products and co-products, and process 

integration. Future work entails obtaining better gas clean up costs for various cleaning 

and conditioning configurations that will be the most beneficial for downstream 

conversion of biomass derived synthesis gas. Additional capital cost items where vendor 

information will reduce the amount of uncertainty in this analysis include a steam 

reformer cost for reforming synthesis gas streams particularly those containing low 

amounts of methane and a PSA cost for gas streams containing less than 70 mol% 

hydrogen. Although the capital cost information for the feed handling and gasification 

come from studies that have used detailed design information, specific breakdowns of the 

cost components as well as operating costs would improve the accuracy of the analysis. 

Another item that should be examined in the future from an environmental point of view 

as well as an economical point of view is flue gas dryers versus steam dryers. More work 

also needs to be done to compare indirect gasification with direct gasification to 

determine the most suitable and economically viable gasification system for different 

fuels products. Future work will also entail examining other biomass feedstocks and 

other products along with the integration of thermochemical and biochemical conversion 

processes into biorefinery concepts. 

54
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Montague, L.; Slayton, A.; Lukas, J. (2002). Lignocellulosic Biomass to Ethanol Process 

Design and Economics Utilizing Co-Current Dilute Acid Prehydrolysis and Enzymatic 

Hydrolysis for Corn Stover. 154 pp.; NREL Report No. TP-510-32438. 

Bain, R. (2004). Personal correspondence. National Renewable Energy Laboratory. 

Golden, CO. 

Bain, R. (January 14, 1992). Material and Energy Balances for Methanol from Biomass 

Using Biomass Gasifiers. 

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www.salaryexpert.com, 2004. 

Breault, R.; Morgan, D. (1992). Design and Economics of Electricity Production Form 

An Indirectly-heated Biomass Gasifier. Report TR4533-049-92. Columbus, Ohio: 

Battelle Columbus Laboratory. 

Chem Systems. (1994). Biomass to Ethanol Process Evaluation. Prepared for NREL. 

Tarrytown, New York. 

Chemical Marketing Reporter. (2004). August 23-30 issues. 

Craig, K.R.; Mann, M.K. (1996). Cost and Performance Analysis of Biomass-Based 

Integrated Gasification Combined-Cycle (BIGCC) Power Systems, National Renewable 

Energy Laboratory, NREL/TP-430-21657. 

Craig, K. (1994). Electric Power Generation Cost - Version 1.11 spreadsheet from 

Kevin Craig. 7/6/94. 

Dravo Engineering Companies. (1987). Gasification Capital Cost Estimation. Obtained 

from Mark Paisley, August, 1994. Battelle Columbus Laboratory. 

East Bay municipal utility district (2004) 

http://www.ebmud.com/wastewater/industrial_&_commercial_permits_&_fees/wastewat 

er_rates/default.htm#non-residential%20rates 

Energy Information Agency (EIA). (2003). 

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monthly/current/pdf/pmmtab16.pdf. 

Feldmann, H. F.; Paisley, M. A.; Applebaus, H.R.; Taylor, D.R. (July 1988). 

Conversion of Forest Residues to A Medium-Rich Gas in a High-Throughput Gasifier. 

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Fogler, H.S. (1992). Elements of Chemical Reaction Engineering. Second Edition. 

Prentice Hall. Englewodd Cliffs, New Jersey. 

Garrett, D.E. (1989). Chemical Engineering Economics. Van Nostrand Reinhold. New 

York. 

Gas Processors Suppliers Association. (2004). Engineering Data Book, FPS Version, 

12th ed., Tulsa, OK. 

Graubard, D. (2004). Personal correspondence. Gas Technology Products LLC. 

Schaumburg, IL. 

Haddeland, G.E. (1980). Waste Treatment Costs. SRI Report No. 137. Menlo Park, CA. 

Jeakel, D. (2004). Price quote from AGSCO for super sacks or bulk. 

Kohl, A.L.; Nielsen, R.B. (1997). Gas Purification. Fifth Edition. Gulf Publishing 

Company. 

Leiby, S.M. (1994). Options for Refinery Hydrogen. SRI Report No. 212. Menlo Park, 

CA. 

Mann, M.K.; Spath, P.L. (1997). Life Cycle Assessment of a Biomass Gasification 

Combined-Cycle Power System. National Renewable Energy Laboratory, Golden, CO, 

TP-430-23076. 

Mann, M. K. (1995). Technical and Economic Assessment of Producing Hydrogen by 

Reforming Syngas from the Battelle Indirectly Heated Biomass Gasifier. National 

Renewable Energy Laboratory, NREL/TP-431-8143. 

Overend, R. (2004). Personal correspondence. National Renewable Energy Laboratory. 

Golden, CO. 

Perry, R.H., Green, D.W., Maloney, J.O. (1997). Perry’s Chemical Engineers’ 

Handbook, 7th ed., McGraw-Hill. 

Peters, M.S. and K.D. Timmerhaus. (2003). Plant Design and Economics for Chemical 

Engineers, 5th Edition, McGraw-Hill, Inc., New York. 

Peters, M.S. and K.D. Timmerhaus. (1991). Plant Design and Economics for Chemical 

Engineers, 4th Edition, McGraw-Hill, Inc., New York. 

Phillips, S.; Carpenter, D.; Dayton, D.; Feik, C.; French, R.; Ratcliff, M.; Hansen, R.; 

Deutch, S.; and Michener, B. (2004). Preliminary Report on the Performance of Full 

Stream Tar Reformer. Internal NREL Milestone report. 

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Schendel, R.L.; Mariz, C.L.; Mak, J.Y. (August 1983). Hydrocarbon Processing. 

Volume 62, p. 58. 

Spath, P. L.; Dayton, D. C. (2003). Preliminary Screening -- Technical and Economic 

Assessment of Synthesis Gas to Fuels and Chemicals with Emphasis on the Potential for 

Biomass-Derived Syngas. 160 pp.; NREL Report No. TP-510-34929. 

Spath, P.L., Mann, M.K. (2000). Life Cycle Assessment of a Natural Gas CombinedCycle Power Generation System, National Renewable Energy Laboratory, NREL/TP570-27715. 

SRI International. (2003). PEP Yearbook International. Volume 1E. United States. 

Menlo Park, CA. 

Stone & Webster; Weyerhaeuser; Amoco; and Carolina Power & Light. (June 1995). 

New Bern Biomass to Energy Project Phase 1 Feasibility Study. Response to NREL 

Contract No. LOI No. RCA-3-13326. NREL Report No. TP-421-7942. 

Walas, S.M. (1988) Chemical Process Equipment Selection and Design. ButterworthHeinemann. 

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Gasification Systems," in Energy from Biomass and Wastes XIII, Donald L. Klass ed. 

Chicago: Institute of Gas Technology, pp.791-827. 

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Boca Raton, FL. 

Weyerhaeuser, Nexant, and Stone & Webster. (2000). Biomass Gasification Combined 

Cycle. Weyerhaeuser Company, Tacoma, WA . DOE DE-FC36-96GO10173. 

Weyerhaeuser. (1992). Gasification Capital Cost Estimation. Obtained from Mark 

Paisley, August, 1994. Battelle Columbus Laboratory. 

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9201/01-SQ. 

57
Appendix A: Current and Goal Base Case Summary Sheets 
Hydrogen Production Process Engineering Analysis 

Design Report: Current Case 

2000 Dry Metric Tonnes Biomass per Day 

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle 

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.38 $9.62 ($/GJ H2, HHV basis) 

$11.48 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 42.6

 Gasification, Tar Reforming, & Quench $16,800,000 Natural Gas 2.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.3

 Steam Methane Reforming, Shift, and PSA $30,300,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 7.1

 Steam System and Power Generation $14,200,000 Other Raw Materials 0.9

 Cooling Water and Other Utilities $3,400,000 Waste Disposal 1.3 

Total Installed Equipment Cost $101,700,000 Electricity 7.5 

Fixed Costs 18.7

 Indirect Costs 51,900,000 Capital Depreciation 14.2

 (% of TPI) 33.8% Average Income Tax 11.0 

Average Return on Investment 26.5 

Total Project Investment (TPI) $153,600,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $1,500,000 

Term (years) N/A Tar Cracking Catalyst $2,400,000 

Capital Charge Factor 0.183 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $500,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $700,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $4,100,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $10,200,000 

Capital Depreciation $7,700,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $6,000,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $14,400,000 

Overall Plant Efficiency - HHV 51.0% 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Goal Case 

2000 Dry Metric Tonnes Biomass per Day 

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle 

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.24 $8.66 ($/GJ H2, HHV basis) 

$10.34 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 39.7

 Gasification, Tar Reforming/Regeneration, & Quench $23,800,000 Natural Gas 5.9

 Compression & Sulfur Removal $16,100,000 Catalysts 0.7

 Shift, and PSA $16,500,000 Olivine 6.6

 Hydrogen Compression $2,800,000 Other Raw Materials 0.9

 Steam System and Power Generation $14,200,000 Waste Disposal 1.2

 Cooling Water and Other Utilities $3,400,000 Electricity 7.1 

Total Installed Equipment Cost $95,700,000 Fixed Costs 16.8 

Capital Depreciation 12.3

 Indirect Costs 48,800,000 Average Income Tax 9.8

 (% of TPI) 33.8% Average Return on Investment 23.3 

Total Project Investment (TPI) $144,400,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $3,400,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.184 Other Raw Matl. Costs $500,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $4,100,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $9,800,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,200,000 

Current Yield (Actual/Theoretical) 49% Average Income Tax $5,700,000 

Average Return on Investment $13,600,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 40259 

Overall Plant Efficiency - HHV 53.3% Electricity Produced Onsite (KW) -29974 

Overall Plant Efficiency - LHV 47.8% Electricity Purchased from Grid (KW) 10284 

Plant Electricity Use (KWh/kg H2) 5.79 

Plant Steam Use (kg steam/kg H2) 19.5 
Appendix B: Sensitivity Summary Sheets 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - $0 feed cost (Case A)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $0.94 $6.58 ($/GJ H2, HHV basis) 

$7.86 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $0 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 0.0

 Gasification, Tar Reforming, & Quench $16,800,000 Natural Gas 2.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.3

 Steam Methane Reforming, Shift, and PSA $30,300,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 7.1

 Steam System and Power Generation $14,200,000 Other Raw Materials 0.8

 Cooling Water and Other Utilities $3,400,000 Waste Disposal 1.3 

Total Installed Equipment Cost $101,700,000 Electricity 7.5 

Fixed Costs 18.7

 Indirect Costs 51,900,000 Capital Depreciation 14.2

 (% of TPI) 33.8% Average Income Tax 11.0 

Average Return on Investment 25.9 

Total Project Investment (TPI) $153,600,000 

Operating Costs ($/yr) 

Feedstock $0 

Loan Rate N/A Natural Gas $1,500,000 

Term (years) N/A Tar Cracking Catalyst $2,400,000 

Capital Charge Factor 0.181 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $700,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $4,100,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $10,200,000 

Capital Depreciation $7,700,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $6,000,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $14,100,000 

Overall Plant Efficiency - HHV 51.0% 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - $53/dry ton feed cost (Case B)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.71 $11.92 ($/GJ H2, HHV basis) 

$14.24 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $53 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

$18,900,000 

$16,800,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$101,700,000 

Operating Costs (cents/kg hydrogen)

Feedstock 

Natural Gas 

Tar Reforming Catalyst 

Other Catalysts 

Olivine 

Other Raw Materials 

Waste Disposal 

Electricity 

Fixed Costs 

75.3

2.8

4.3

0.7

7.1

0.8

1.3 

7.5 

18.7

 Indirect Costs 

(% of TPI) 

Total Project Investment (TPI) 

51,900,000 

33.8% 

$153,600,000 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Operating Costs ($/yr) 

Feedstock 

14.2

11.1 

26.9 

$40,900,000 

Loan Rate 

Term (years) 

Capital Charge Factor 

N/A 

N/A 

0.184 

Natural Gas 

Tar Cracking Catalyst 

Other Catalysts 

Olivine 

$1,500,000 

$2,400,000 

$400,000 

$3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

119.7 

155.0 

45% 

Waste Disposal 

Electricity 

Fixed Costs 

$700,000

$4,100,000 

$10,200,000 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

72.14% 

71.78% 

51.0% 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

$7,700,000 

$6,000,000 

$14,600,000 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

35803

-25583

10219 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

5.54 

23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - 30% moisture feedstock (Case C)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.31 $9.12 ($/GJ H2, HHV basis) 

$10.89 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 66 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 907 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $15,600,000 Feedstock 42.6

 Gasification, Tar Reforming, & Quench $20,000,000 Natural Gas 2.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.3

 Steam Methane Reforming, Shift, and PSA $30,400,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 7.1

 Steam System and Power Generation $16,700,000 Other Raw Materials 1.0

 Cooling Water and Other Utilities $3,500,000 Waste Disposal 1.3 

Total Installed Equipment Cost $104,300,000 Electricity -1.1 

Fixed Costs 18.8

 Indirect Costs 53,200,000 Capital Depreciation 14.5

 (% of TPI) 33.8% Average Income Tax 11.5 

Average Return on Investment 27.0 

Total Project Investment (TPI) $157,500,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $300,000 

Term (years) N/A Tar Cracking Catalyst $2,400,000 

Capital Charge Factor 0.183 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $1,700,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $700,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity -$600,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $10,200,000 

Capital Depreciation $7,900,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $6,300,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $14,700,000 

Overall Plant Efficiency - HHV 54.7% 

Overall Plant Efficiency - LHV 49.5% Total Plant Electricity Usage (KW) 36697

 Electricity Produced Onsite (KW) -38226

 Electricity Purchased from Grid (KW) -1529 

Plant Electricity Use (KWh/kg H2) 5.67 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - dry to 20% moisture (lower gasifier temp) (Case D)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.37 $9.57 ($/GJ H2, HHV basis) 

$11.44 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 52.4 63.4 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 67.9 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$12,900,000 

$17,800,000 

$15,300,000 

$29,700,000 

$2,600,000 

$14,700,000 

$3,700,000 

$96,700,000 

49,300,000 

33.8% 

$145,900,000 

N/A 

N/A 

0.184 

119.7 

155.0 

44% 

70.14% 

69.83% 

50.2% 

45.0% 

Operating Costs (cents/kg hydrogen)

Feedstock 44.2

Natural Gas 2.9

Tar Reforming Catalyst 4.9

Other Catalysts 0.7

Olivine 7.3

Other Raw Materials 0.9

Waste Disposal 1.5 

Electricity 4.8 

Fixed Costs 18.8

Capital Depreciation 13.9

Average Income Tax 11.0 

Average Return on Investment 26.1 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $1,500,000 

Tar Cracking Catalyst $2,500,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $500,000

Waste Disposal $800,000

Electricity $2,500,000 

Fixed Costs $9,900,000 

Capital Depreciation $7,300,000 

Average Income Tax $5,800,000 

Average Return on Investment $13,700,000 

Total Plant Electricity Usage (KW) 35854

 Electricity Produced Onsite (KW) -29525

 Electricity Purchased from Grid (KW) 6328 

Plant Electricity Use (KWh/kg H2) 5.75 

Plant Steam Use (kg steam/kg H2) 23.4 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - dry to 20% moisture with same gasifier temperature (Case E) 

2000 Dry Metric Tonnes Biomass per Day 

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle 

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg)$1.39 $9.70 ($/GJ H2, HHV basis) 

$11.59 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 53.8 65.1 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 69.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $12,900,000 Feedstock 43.0

 Gasification, Tar Reforming, & Quench $19,000,000 Natural Gas 5.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.8

 Steam Methane Reforming, Shift, and PSA $30,200,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 8.0

 Steam System and Power Generation $15,000,000 Other Raw Materials 0.9

 Cooling Water and Other Utilities $3,700,000 Waste Disposal 1.6 

Total Installed Equipment Cost $98,900,000 Electricity 4.6 

Fixed Costs 18.6

 Indirect Costs 50,500,000 Capital Depreciation 13.9 

(% of TPI) 33.8% Average Income Tax 11.0 

Average Return on Investment 26.0 

Total Project Investment (TPI) $149,400,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $3,100,000 

Term (years) N/A Tar Cracking Catalyst $2,600,000 

Capital Charge Factor 0.183 Other Catalysts $400,000 

Olivine $4,300,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $500,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $800,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $2,500,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $10,000,000 

Capital Depreciation $7,500,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $5,900,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $14,000,000 

Overall Plant Efficiency - HHV 50.4% 

Overall Plant Efficiency - LHV 45.3% Total Plant Electricity Usage (KW) 36588

 Electricity Produced Onsite (KW) -30442

 Electricity Purchased from Grid (KW) 6146 

Plant Electricity Use (KWh/kg H2) 5.71 

Plant Steam Use (kg steam/kg H2) 23.2 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - dry to 30% moisture (lower gasifier temp) (Case F)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.46 $10.22 ($/GJ H2, HHV basis) 

$12.20 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 48.8 59.0 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 63.2 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $10,400,000 Feedstock 47.5

 Gasification, Tar Reforming, & Quench $20,100,000 Natural Gas 9.5

 Compression & Sulfur Removal $15,200,000 Tar Reforming Catalyst 5.8

 Steam Methane Reforming, Shift, and PSA $28,700,000 Other Catalysts 0.8

 Hydrogen Compression $2,400,000 Olivine 7.9

 Steam System and Power Generation $16,900,000 Other Raw Materials 1.3

 Cooling Water and Other Utilities $4,200,000 Waste Disposal 1.8 

Total Installed Equipment Cost $97,900,000 Electricity -4.4 

Fixed Costs 20.4

 Indirect Costs 49,900,000 Capital Depreciation 15.2

 (% of TPI) 33.8% Average Income Tax 12.2 

Average Return on Investment 28.4 

Total Project Investment (TPI) $147,700,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $4,600,000 

Term (years) N/A Tar Cracking Catalyst $2,800,000 

Capital Charge Factor 0.185 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $600,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $900,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity -$2,100,000 

Current Yield (Actual/Theoretical) 41% Fixed Costs $10,000,000 

Capital Depreciation $7,400,000 

Gasifier Efficiency - HHV 67.08% Average Income Tax $6,000,000 

Gasifier Efficiency - LHV 66.84% Average Return on Investment $13,900,000 

Overall Plant Efficiency - HHV 48.9% 

Overall Plant Efficiency - LHV 44.8% Total Plant Electricity Usage (KW) 37588

 Electricity Produced Onsite (KW) -42891

 Electricity Purchased from Grid (KW) -5303 

Plant Electricity Use (KWh/kg H2) 6.47 

Plant Steam Use (kg steam/kg H2) 23.5 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - dry to 30% moisture with same gasifier temperature (Case G)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.50 $10.46 ($/GJ H2, HHV basis) 

$12.50 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 53.0 64.1 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 68.6 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$10,500,000 

$23,300,000 

$15,700,000 

$30,000,000 

$2,600,000 

$17,800,000 

$4,300,000 

$104,200,000 

53,100,000 

33.8% 

$157,300,000 

N/A 

N/A 

0.185 

119.7 

155.0 

44% 

72.02% 

71.66% 

49.8% 

45.8% 

Operating Costs (cents/kg hydrogen)

Feedstock 43.7

Natural Gas 17.8

Tar Reforming Catalyst 5.6

Other Catalysts 0.7

Olivine 9.4

Other Raw Materials 1.3

Waste Disposal 1.9 

Electricity -4.6 

Fixed Costs 19.5

Capital Depreciation 14.9

Average Income Tax 12.0 

Average Return on Investment 27.8 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $9,400,000 

Tar Cracking Catalyst $2,900,000 

Other Catalysts $400,000 

Olivine $5,000,000 

Other Raw Matl. Costs $700,000

Waste Disposal $1,000,000

Electricity -$2,500,000 

Fixed Costs $10,300,000 

Capital Depreciation $7,900,000 

Average Income Tax $6,400,000 

Average Return on Investment $14,800,000 

Total Plant Electricity Usage (KW) 39856

 Electricity Produced Onsite (KW) -46002

 Electricity Purchased from Grid (KW) -6146 

Plant Electricity Use (KWh/kg H2) 6.32 

Plant Steam Use (kg steam/kg H2) 22.8 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - no dryer (Case H)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.78 $12.43 ($/GJ H2, HHV basis) 

$14.85 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 51.3 62 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 66.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $8,200,000 Feedstock 45.2

 Gasification, Tar Reforming, & Quench $31,900,000 Natural Gas 45.9

 Compression & Sulfur Removal $16,200,000 Tar Reforming Catalyst 8.0

 Steam Methane Reforming, Shift, and PSA $29,600,000 Other Catalysts 0.7

 Hydrogen Compression $2,500,000 Olivine 14.2

 Steam System and Power Generation $23,500,000 Other Raw Materials 2.2

 Cooling Water and Other Utilities $5,600,000 Waste Disposal 3.1 

Total Installed Equipment Cost $117,500,000 Electricity -26.8 

Fixed Costs 21.5

 Indirect Costs 59,900,000 Capital Depreciation 17.4

 (% of TPI) 33.8% Average Income Tax 14.4 

Average Return on Investment 32.4 

Total Project Investment (TPI) $177,200,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $900,000 

Term (years) N/A Tar Cracking Catalyst $4,100,000 

Capital Charge Factor 0.186 Other Catalysts $400,000 

Olivine $7,300,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $23,800,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $1,600,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity -$13,800,000 

Current Yield (Actual/Theoretical) 43% Fixed Costs $11,000,000 

Capital Depreciation $8,900,000 

Gasifier Efficiency - HHV 72.01% Average Income Tax $7,400,000 

Gasifier Efficiency - LHV 71.65% Average Return on Investment $16,600,000 

Overall Plant Efficiency - HHV 52.0% 

Overall Plant Efficiency - LHV 50.6% Total Plant Electricity Usage (KW) 46376

 Electricity Produced Onsite (KW) -80705

 Electricity Purchased from Grid (KW) -34329 

Plant Electricity Use (KWh/kg H2) 7.60 

Plant Steam Use (kg steam/kg H2) 21.8 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - stm:wood ratio = 0.1 (Case I)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.30 $9.10 ($/GJ H2, HHV basis) 

$10.87 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 56.3 68 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 72.9 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$19,300,000 

$16,700,000 

$13,800,000 

$30,900,000 

$2,700,000 

$13,500,000 

$3,400,000 

$100,300,000 

51,100,000 

33.8% 

$151,400,000 

N/A 

N/A 

0.183 

119.7 

155.0 

47% 

72.05% 

71.69% 

52.5% 

46.9% 

Operating Costs (cents/kg hydrogen)

Feedstock 41.1

Natural Gas 2.6

Tar Reforming Catalyst 3.1

Other Catalysts 0.7

Olivine 5.9

Other Raw Materials 0.7

Waste Disposal 1.1 

Electricity 8.4 

Fixed Costs 17.7

Capital Depreciation 13.5

Average Income Tax 10.4 

Average Return on Investment 25.2 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $200,000 

Tar Cracking Catalyst $1,800,000 

Other Catalysts $400,000 

Olivine $3,300,000 

Other Raw Matl. Costs $1,600,000

Waste Disposal $600,000

Electricity $4,700,000 

Fixed Costs $10,000,000 

Capital Depreciation $7,600,000 

Average Income Tax $5,900,000 

Average Return on Investment $14,200,000 

Total Plant Electricity Usage (KW) 34388

 Electricity Produced Onsite (KW) -22657

 Electricity Purchased from Grid (KW) 11732 

Plant Electricity Use (KWh/kg H2) 5.13 

Plant Steam Use (kg steam/kg H2) 16.9 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - stm:wood ratio = 1 & lower gasifier temp (Case J)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.57 $10.94 ($/GJ H2, HHV basis) 

$13.07 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 48.6 58.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 62.9 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $19,300,000 Feedstock 47.7

 Gasification, Tar Reforming, & Quench $18,400,000 Natural Gas 6.5

 Compression & Sulfur Removal $15,600,000 Tar Reforming Catalyst 7.2

 Steam Methane Reforming, Shift, and PSA $28,500,000 Other Catalysts 0.8

 Hydrogen Compression $2,400,000 Olivine 7.9

 Steam System and Power Generation $16,400,000 Other Raw Materials 1.4

 Cooling Water and Other Utilities $4,000,000 Waste Disposal 2.1 

Total Installed Equipment Cost $104,600,000 Electricity 2.2 

Fixed Costs 21.3

 Indirect Costs 53,300,000 Capital Depreciation 16.3

 (% of TPI) 33.8% Average Income Tax 12.9 

Average Return on Investment 30.5 

Total Project Investment (TPI) $157,900,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $3,100,000 

Term (years) N/A Tar Cracking Catalyst $3,500,000 

Capital Charge Factor 0.184 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $700,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $1,000,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $1,100,000 

Current Yield (Actual/Theoretical) 41% Fixed Costs $10,400,000 

Capital Depreciation $7,900,000 

Gasifier Efficiency - HHV 68.06% Average Income Tax $6,300,000 

Gasifier Efficiency - LHV 67.79% Average Return on Investment $14,800,000 

Overall Plant Efficiency - HHV 46.3% 

Overall Plant Efficiency - LHV 41.7% Total Plant Electricity Usage (KW) 37246

 Electricity Produced Onsite (KW) -34550

 Electricity Purchased from Grid (KW) 2696 

Plant Electricity Use (KWh/kg H2) 6.45 

Plant Steam Use (kg steam/kg H2) 42.0 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - stm:wood ratio = 1 with same gasifier temperature (Case K)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.58 $11.04 ($/GJ H2, HHV basis) 

$13.19 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 52.1 63 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 67.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $19,300,000 Feedstock 44.5

 Gasification, Tar Reforming, & Quench $20,900,000 Natural Gas 12.7

 Compression & Sulfur Removal $16,000,000 Tar Reforming Catalyst 6.9

 Steam Methane Reforming, Shift, and PSA $29,700,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 9.2

 Steam System and Power Generation $17,000,000 Other Raw Materials 1.4

 Cooling Water and Other Utilities $4,000,000 Waste Disposal 2.2 

Total Installed Equipment Cost $109,500,000 Electricity 2.0 

Fixed Costs 20.2

 Indirect Costs 55,800,000 Capital Depreciation 15.9

 (% of TPI) 33.8% Average Income Tax 12.6 

Average Return on Investment 29.7 

Total Project Investment (TPI) $165,300,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $400,000 

Term (years) N/A Tar Cracking Catalyst $3,600,000 

Capital Charge Factor 0.184 Other Catalysts $400,000 

Olivine $4,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $6,900,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $1,100,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $1,100,000 

Current Yield (Actual/Theoretical) 43% Fixed Costs $10,500,000 

Capital Depreciation $8,300,000 

Gasifier Efficiency - HHV 72.10% Average Income Tax $6,600,000 

Gasifier Efficiency - LHV 71.74% Average Return on Investment $15,500,000 

Overall Plant Efficiency - HHV 47.2% 

Overall Plant Efficiency - LHV 42.6% Total Plant Electricity Usage (KW) 38893

 Electricity Produced Onsite (KW) -36241

 Electricity Purchased from Grid (KW) 2651 

Plant Electricity Use (KWh/kg H2) 6.28 

Plant Steam Use (kg steam/kg H2) 40.1 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - no H2 recycle (Case L)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.30 $9.10 ($/GJ H2, HHV basis) 

$10.87 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.6 71 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.9 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,800,000 

$15,500,000 

$30,800,000 

$2,800,000 

$13,000,000 

$3,200,000 

$101,000,000 

51,500,000 

33.8% 

$152,400,000 

N/A 

N/A 

0.183 

119.7 

155.0 

49% 

72.14% 

71.78% 

53.6% 

47.7% 

Operating Costs (cents/kg hydrogen)

Feedstock 39.5

Natural Gas 2.2

Tar Reforming Catalyst 4.0

Other Catalysts 0.7

Olivine 6.6

Other Raw Materials 0.6

Waste Disposal 1.2 

Electricity 10.9 

Fixed Costs 17.1

Capital Depreciation 13.0

Average Income Tax 10.0 

Average Return on Investment 24.5 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $200,000 

Tar Cracking Catalyst $2,400,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $1,400,000

Waste Disposal $700,000

Electricity $6,400,000 

Fixed Costs $10,000,000 

Capital Depreciation $7,600,000 

Average Income Tax $5,900,000 

Average Return on Investment $14,400,000 

Total Plant Electricity Usage (KW) 35666

 Electricity Produced Onsite (KW) -19721

 Electricity Purchased from Grid (KW) 15944 

Plant Electricity Use (KWh/kg H2) 5.12 

Plant Steam Use (kg steam/kg H2) 21.9 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - no LTS (Case M)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.47 $10.24 ($/GJ H2, HHV basis) 

$12.23 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 49.3 60 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 63.9 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,900,000 

$15,500,000 

$29,500,000 

$2,500,000 

$15,500,000 

$3,600,000 

$102,400,000 

52,200,000 

33.8% 

$154,500,000 

N/A 

N/A 

0.183 

119.7 

155.0 

41% 

72.14% 

71.78% 

47.9% 

43.0% 

Operating Costs (cents/kg hydrogen)

Feedstock 46.9

Natural Gas 3.1

Tar Reforming Catalyst 4.8

Other Catalysts 0.8

Olivine 7.8

Other Raw Materials 1.0

Waste Disposal 1.5 

Electricity 3.1 

Fixed Costs 20.5

Capital Depreciation 15.6

Average Income Tax 12.4 

Average Return on Investment 29.4 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $300,000 

Tar Cracking Catalyst $2,400,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $1,700,000

Waste Disposal $700,000

Electricity $1,500,000 

Fixed Costs $10,100,000 

Capital Depreciation $7,700,000 

Average Income Tax $6,100,000 

Average Return on Investment $14,500,000 

Total Plant Electricity Usage (KW) 35941

 Electricity Produced Onsite (KW) -32124

 Electricity Purchased from Grid (KW) 3817 

Plant Electricity Use (KWh/kg H2) 6.12 

Plant Steam Use (kg steam/kg H2) 26.0 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - 0.5% tar reformer catalyst loss (Case N)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.35 $9.43 ($/GJ H2, HHV basis) 

$11.27 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 66 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,800,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$101,700,000 

51,900,000 

33.8% 

$153,600,000 

N/A 

N/A 

0.183 

119.7 

155.0 

45% 

72.14% 

71.78% 

51.0% 

45.6% 

Operating Costs (cents/kg hydrogen)

Feedstock 42.6

Natural Gas 2.8

Tar Reforming Catalyst 2.2

Other Catalysts 0.7

Olivine 7.1

Other Raw Materials 0.8

Waste Disposal 1.3 

Electricity 7.5 

Fixed Costs 18.5

Capital Depreciation 14.2

Average Income Tax 11.0 

Average Return on Investment 26.4 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $200,000 

Tar Cracking Catalyst $1,200,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $1,700,000

Waste Disposal $700,000

Electricity $4,100,000 

Fixed Costs $10,100,000 

Capital Depreciation $7,700,000 

Average Income Tax $6,000,000 

Average Return on Investment $14,400,000 

Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Current Case - Internal waste water treatment (Case O)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.38 $9.62 ($/GJ H2, HHV basis) 

$11.49 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 66 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,800,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$101,700,000 

52,500,000 

33.8% 

$155,500,000 

N/A 

N/A 

0.183 

119.7 

155.0 

45% 

72.14% 

71.78% 

51.0% 

45.7% 

Operating Costs (cents/kg hydrogen)

Feedstock 42.6

Natural Gas 2.8

Tar Reforming Catalyst 4.3

Other Catalysts 0.7

Olivine 7.1

Other Raw Materials 0.7

Waste Disposal 1.2 

Electricity 7.4 

Fixed Costs 18.7

Capital Depreciation 14.3

Average Income Tax 11.1 

Average Return on Investment 26.8 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $200,000 

Tar Cracking Catalyst $2,400,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $1,700,000

Waste Disposal $700,000

Electricity $4,000,000 

Fixed Costs $10,100,000 

Capital Depreciation $7,800,000 

Average Income Tax $6,100,000 

Average Return on Investment $14,600,000 

Total Plant Electricity Usage (KW) 35814

 Electricity Produced Onsite (KW) -25752

 Electricity Purchased from Grid (KW) 10063 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sesitivity on Current Case - Increase PSA cost (Case P)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.42 $9.90 ($/GJ H2, HHV basis) 

$11.82 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

$18,900,000 

$16,800,000 

$15,500,000 

$37,500,000 

$2,600,000 

$14,200,000 

$3,400,000 

$108,900,000 

Operating Costs (cents/kg hydrogen)

Feedstock 

Natural Gas 

Tar Reforming Catalyst 

Other Catalysts 

Olivine 

Other Raw Materials 

Waste Disposal 

Electricity 

Fixed Costs 

42.6

2.8

4.3

0.7

7.1

0.8

1.3 

7.5 

19.5

 Indirect Costs 

(% of TPI) 

Total Project Investment (TPI) 

55,600,000 

33.8% 

$164,400,000 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Operating Costs ($/yr) 

Feedstock 

15.1

11.7 

28.3 

$23,200,000 

Loan Rate 

Term (years) 

Capital Charge Factor 

N/A 

N/A 

0.182 

Natural Gas 

Tar Cracking Catalyst 

Other Catalysts 

Olivine 

$1,500,000 

$2,400,000 

$400,000 

$3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

119.7 

155.0 

45% 

Waste Disposal 

Electricity 

Fixed Costs 

$700,000

$4,100,000 

$10,600,000 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

72.14% 

71.78% 

51.0% 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

$8,200,000 

$6,400,000 

$15,400,000 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

35803

-25583

10219 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

5.54 

23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - Increase in steam reformer cost (Case Q)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.45 $10.11 ($/GJ H2, HHV basis) 

$12.07 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,800,000 

$15,500,000 

$42,600,000 

$2,600,000 

$14,200,000 

$3,400,000 

$114,000,000 

58,100,000 

33.8% 

$172,100,000 

N/A 

N/A 

0.182 

119.7 

155.0 

45% 

72.14% 

71.78% 

51.0% 

45.6% 

Operating Costs (cents/kg hydrogen)

Feedstock 42.6

Natural Gas 2.8

Tar Reforming Catalyst 4.3

Other Catalysts 0.7

Olivine 7.1

Other Raw Materials 0.8

Waste Disposal 1.3 

Electricity 7.5 

Fixed Costs 20.1

Capital Depreciation 15.8

Average Income Tax 12.1 

Average Return on Investment 29.5 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $1,500,000 

Tar Cracking Catalyst $2,400,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $400,000

Waste Disposal $700,000

Electricity $4,100,000 

Fixed Costs $10,900,000 

Capital Depreciation $8,600,000 

Average Income Tax $6,600,000 

Average Return on Investment $16,100,000 

Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - Higher Electricity Cost (Case R)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.40 $9.75 ($/GJ H2, HHV basis) 

$11.64 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,800,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$101,700,000 

51,900,000 

33.8% 

$153,600,000 

N/A 

N/A 

0.183 

119.7 

155.0 

45% 

72.14% 

71.78% 

51.0% 

45.6% 

Operating Costs (cents/kg hydrogen)

Feedstock 42.6

Natural Gas 2.8

Tar Reforming Catalyst 4.3

Other Catalysts 0.7

Olivine 7.1

Other Raw Materials 0.8

Waste Disposal 1.3 

Electricity 9.5 

Fixed Costs 18.7

Capital Depreciation 14.2

Average Income Tax 11.0 

Average Return on Investment 26.5 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $1,500,000 

Tar Cracking Catalyst $2,400,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $400,000

Waste Disposal $700,000

Electricity $5,200,000 

Fixed Costs $10,200,000 

Capital Depreciation $7,700,000 

Average Income Tax $6,000,000 

Average Return on Investment $14,400,000 

Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - Higher Natural Gas Cost (Case S)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.39 $9.67 ($/GJ H2, HHV basis) 

$11.55 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

$18,900,000 

$16,800,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$101,700,000 

51,900,000 

33.8% 

$153,600,000 

N/A 

N/A 

0.183 

119.7 

155.0 

45% 

72.14% 

71.78% 

51.0% 

45.6% 

Operating Costs (cents/kg hydrogen)

Feedstock 42.6

Natural Gas 3.7

Tar Reforming Catalyst 4.3

Other Catalysts 0.7

Olivine 7.1

Other Raw Materials 0.8

Waste Disposal 1.3 

Electricity 7.5 

Fixed Costs 18.7

Capital Depreciation 14.2

Average Income Tax 11.0 

Average Return on Investment 26.5 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $2,000,000 

Tar Cracking Catalyst $2,400,000 

Other Catalysts $400,000 

Olivine $3,800,000 

Other Raw Matl. Costs $400,000

Waste Disposal $700,000

Electricity $4,100,000 

Fixed Costs $10,200,000 

Capital Depreciation $7,700,000 

Average Income Tax $6,000,000 

Average Return on Investment $14,400,000 

Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - Low Feed Handling & Drying Cost (Case T)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.35 $9.41 ($/GJ H2, HHV basis) 

$11.24 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

$14,200,000 

$16,800,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$97,000,000 

Operating Costs (cents/kg hydrogen)

Feedstock 

Natural Gas 

Tar Reforming Catalyst 

Other Catalysts 

Olivine 

Other Raw Materials 

Waste Disposal 

Electricity 

Fixed Costs 

42.6

2.8

4.3

0.7

7.1

0.8

1.3 

7.5 

18.2

 Indirect Costs 

(% of TPI) 

Total Project Investment (TPI) 

49,500,000 

33.8% 

$146,400,000 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Operating Costs ($/yr) 

Feedstock 

13.4

10.6 

25.4 

$23,200,000 

Loan Rate 

Term (years) 

Capital Charge Factor 

N/A 

N/A 

0.184 

Natural Gas 

Tar Cracking Catalyst 

Other Catalysts 

Olivine 

$1,500,000 

$2,400,000 

$400,000 

$3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

119.7 

155.0 

45% 

Waste Disposal 

Electricity 

Fixed Costs 

$700,000

$4,100,000 

$9,900,000 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

72.14% 

71.78% 

51.0% 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

$7,300,000 

$5,800,000 

$13,800,000 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

35803

-25583

10219 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

5.54 

23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - Low Gasification & Clean Up Cost (Case U)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.35 $9.40 ($/GJ H2, HHV basis) 

$11.22 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

$18,900,000 

$11,700,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$96,600,000 

Operating Costs (cents/kg hydrogen)

Feedstock 

Natural Gas 

Tar Reforming Catalyst 

Other Catalysts 

Olivine 

Other Raw Materials 

Waste Disposal 

Electricity 

Fixed Costs 

42.6

2.8

4.3

0.7

7.1

0.8

1.3 

7.5 

18.2

 Indirect Costs 

(% of TPI) 

Total Project Investment (TPI) 

49,300,000 

33.8% 

$145,900,000 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Operating Costs ($/yr) 

Feedstock 

13.4

10.6 

25.2 

$23,200,000 

Loan Rate 

Term (years) 

Capital Charge Factor 

N/A 

N/A 

0.184 

Natural Gas 

Tar Cracking Catalyst 

Other Catalysts 

Olivine 

$1,500,000 

$2,400,000 

$400,000 

$3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

119.7 

155.0 

45% 

Waste Disposal 

Electricity 

Fixed Costs 

$700,000

$4,100,000 

$9,900,000 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

72.14% 

71.78% 

51.0% 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

$7,300,000 

$5,800,000 

$13,700,000 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

35803

-25583

10219 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

5.54 

23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - Low Feed Handling & Drying Cost Combined with Low Gasification & Clean Up Cost (Case V) 

2000 Dry Metric Tonnes Biomass per Day 

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle 

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg)$1.32 $9.20 ($/GJ H2, HHV basis) 

$10.99 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $14,200,000 Feedstock 42.6

 Gasification, Tar Reforming, & Quench $11,700,000 Natural Gas 2.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.3

 Steam Methane Reforming, Shift, and PSA $30,300,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 7.1

 Steam System and Power Generation $14,200,000 Other Raw Materials 0.8

 Cooling Water and Other Utilities $3,400,000 Waste Disposal 1.3 

Total Installed Equipment Cost $91,900,000 Electricity 7.5 

Fixed Costs 17.6

 Indirect Costs 46,900,000 Capital Depreciation 12.7 

(% of TPI) 33.8% Average Income Tax 10.1 

Average Return on Investment 24.2 

Total Project Investment (TPI) $138,700,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $1,500,000 

Term (years) N/A Tar Cracking Catalyst $2,400,000 

Capital Charge Factor 0.185 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $700,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $4,100,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $9,600,000 

Capital Depreciation $6,900,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $5,500,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $13,200,000 

Overall Plant Efficiency - HHV 51.0% 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - High Feed Handling & Drying Cost (Case W) 

2000 Dry Metric Tonnes Biomass per Day 

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle 

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg)$1.41 $9.86 ($/GJ H2, HHV basis) 

$11.78 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $25,100,000 Feedstock 42.6

 Gasification, Tar Reforming, & Quench $16,800,000 Natural Gas 2.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.3

 Steam Methane Reforming, Shift, and PSA $30,300,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 7.1

 Steam System and Power Generation $14,200,000 Other Raw Materials 0.8

 Cooling Water and Other Utilities $3,400,000 Waste Disposal 1.3 

Total Installed Equipment Cost $107,900,000 Electricity 7.5 

Fixed Costs 19.4

 Indirect Costs 55,100,000 Capital Depreciation 14.9 

(% of TPI) 33.8% Average Income Tax 11.6 

Average Return on Investment 28.1 

Total Project Investment (TPI) $163,000,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $1,500,000 

Term (years) N/A Tar Cracking Catalyst $2,400,000 

Capital Charge Factor 0.182 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $700,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $4,100,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $10,600,000 

Capital Depreciation $8,100,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $6,300,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $15,300,000 

Overall Plant Efficiency - HHV 51.0% 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - High Gasification & Clean Up Cost (Case X) 

2000 Dry Metric Tonnes Biomass per Day 

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle 

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg)$1.42 $9.92 ($/GJ H2, HHV basis) 

$11.85 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 42.6

 Gasification, Tar Reforming, & Quench $24,500,000 Natural Gas 2.8

 Compression & Sulfur Removal $15,500,000 Tar Reforming Catalyst 4.3

 Steam Methane Reforming, Shift, and PSA $30,300,000 Other Catalysts 0.7

 Hydrogen Compression $2,600,000 Olivine 7.1

 Steam System and Power Generation $14,200,000 Other Raw Materials 0.8

 Cooling Water and Other Utilities $3,400,000 Waste Disposal 1.3 

Total Installed Equipment Cost $109,400,000 Electricity 7.5 

Fixed Costs 19.6

 Indirect Costs 55,800,000 Capital Depreciation 15.3 

(% of TPI) 33.8% Average Income Tax 11.7 

Average Return on Investment 28.3 

Total Project Investment (TPI) $165,200,000 

Operating Costs ($/yr) 

Feedstock $23,200,000 

Loan Rate N/A Natural Gas $1,500,000 

Term (years) N/A Tar Cracking Catalyst $2,400,000 

Capital Charge Factor 0.182 Other Catalysts $400,000 

Olivine $3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Waste Disposal $700,000

 Theoretical Yield (kg/dry ton) 155.0 Electricity $4,100,000 

Current Yield (Actual/Theoretical) 45% Fixed Costs $10,700,000 

Capital Depreciation $8,300,000 

Gasifier Efficiency - HHV 72.14% Average Income Tax $6,400,000 

Gasifier Efficiency - LHV 71.78% Average Return on Investment $15,400,000 

Overall Plant Efficiency - HHV 51.0% 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 35803

 Electricity Produced Onsite (KW) -25583

 Electricity Purchased from Grid (KW) 10219 

Plant Electricity Use (KWh/kg H2) 5.54 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Current Case - High Feed Handling & Drying Cost Combined with High Gasification & Clean Up Cost (Case Y)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, Methane Reformer, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.46 $10.17 ($/GJ H2, HHV basis) 

$12.15 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 54.4 65.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 70.4 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying 

Gasification, Tar Reforming, & Quench 

Compression & Sulfur Removal 

Steam Methane Reforming, Shift, and PSA 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

$25,100,000 

$24,500,000 

$15,500,000 

$30,300,000 

$2,600,000 

$14,200,000 

$3,400,000 

$115,600,000 

Operating Costs (cents/kg hydrogen)

Feedstock 

Natural Gas 

Tar Reforming Catalyst 

Other Catalysts 

Olivine 

Other Raw Materials 

Waste Disposal 

Electricity 

Fixed Costs 

42.6

2.8

4.3

0.7

7.1

0.8

1.3 

7.5 

20.3

 Indirect Costs 

(% of TPI) 

Total Project Investment (TPI) 

59,000,000 

33.8% 

$174,600,000 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Operating Costs ($/yr) 

Feedstock 

16.0

12.3 

30.0 

$23,200,000 

Loan Rate 

Term (years) 

Capital Charge Factor 

N/A 

N/A 

0.182 

Natural Gas 

Tar Cracking Catalyst 

Other Catalysts 

Olivine 

$1,500,000 

$2,400,000 

$400,000 

$3,800,000 

Maximum Yields (100% of Theoretical) based on composition Other Raw Matl. Costs $400,000

 Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

119.7 

155.0 

45% 

Waste Disposal 

Electricity 

Fixed Costs 

$700,000

$4,100,000 

$11,000,000 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

72.14% 

71.78% 

51.0% 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

$8,700,000 

$6,700,000 

$16,300,000 

Overall Plant Efficiency - LHV 45.6% Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

35803

-25583

10219 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

5.54 

23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitiivity on Goal Case - $0 feed cost (Case AA)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $0.84 $5.84 ($/GJ H2, HHV basis) 

$6.97 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $0 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying $18,900,000 

Gasification, Tar Reforming, & Quench $17,600,000 

Compression & Sulfur Removal $16,100,000 

Tar Reforming Catalyst Regeneration, Shift, and PSA $22,600,000 

Operating Costs (cents/kg hydrogen)

Feedstock 0.0

Natural Gas 5.9

Catalysts 0.6

Olivine 6.6

Other Raw Materials 0.7

Waste Disposal 1.2

Electricity 7.1 

Fixed Costs 16.8 

Capital Depreciation 12.3

Average Income Tax 9.7

Average Return on Investment 22.8 

Operating Costs ($/yr) 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

$2,800,000 

$14,200,000 

$3,400,000 

$95,600,000 

48,800,000 

33.8% 

$144,400,000 

N/A 

N/A 

0.181 

Feedstock 

Natural Gas 

Catalysts 

Olivine 

Other Raw Matl. Costs 

Waste Disposal 

Electricity 

Fixed Costs 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

$0 

$3,400,000 

$400,000 

$3,800,000 

$400,000 

$700,000 

$4,100,000

$9,800,000

$7,200,000 

$5,600,000 

$13,300,000 

40259 

-29974 

10284 

5.79 

19.5 

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

119.7 

155.0 

49% 

72.14% 

71.78% 

53.3%

47.8%
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Goal Case- $53/dry ton feed cost (Case BB)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.55 $10.80 ($/GJ H2, HHV basis) 

$12.90 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $53 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying $18,900,000 

Gasification, Tar Reforming, & Quench $17,600,000 

Compression & Sulfur Removal $16,100,000 

Tar Reforming Catalyst Regeneration, Shift, and PSA $22,600,000 

Operating Costs (cents/kg hydrogen)

Feedstock 70.0

Natural Gas 5.9

Catalysts 0.6

Olivine 6.6

Other Raw Materials 0.7

Waste Disposal 1.2

Electricity 7.1 

Fixed Costs 16.8 

Capital Depreciation 12.3

Average Income Tax 9.8

Average Return on Investment 23.8 

Operating Costs ($/yr) 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

$2,800,000 

$14,200,000 

$3,400,000 

$95,600,000 

48,800,000 

33.8% 

$144,400,000 

N/A 

N/A 

0.186 

Feedstock 

Natural Gas 

Catalysts 

Olivine 

Other Raw Matl. Costs 

Waste Disposal 

Electricity 

Fixed Costs 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

$40,900,000 

$3,400,000 

$400,000 

$3,800,000 

$400,000 

$700,000 

$4,100,000

$9,800,000

$7,200,000 

$5,700,000 

$13,900,000 

40259 

-29974 

10284 

5.79 

19.5 

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

119.7 

155.0 

49% 

72.14% 

71.78% 

53.3%

47.8%
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Goal Case - 30% moisture feedstock (Case CC)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.18 $8.22 ($/GJ H2, HHV basis) 

$9.81 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 907 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $15,600,000 Feedstock 39.7

 Gasification, Tar Reforming, & Quench $21,100,000 Natural Gas 5.9

 Compression & Sulfur Removal $16,100,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $22,600,000 Olivine 6.6

 Hydrogen Compression $2,800,000 Other Raw Materials 0.9

 Steam System and Power Generation $16,700,000 Waste Disposal 1.2

 Cooling Water and Other Utilities $3,500,000 Electricity -1.0 

Total Installed Equipment Cost $98,400,000 Fixed Costs 17.1 

Capital Depreciation 12.7

 Indirect Costs 50,200,000 Average Income Tax 10.2

 (% of TPI) 33.8% Average Return on Investment 23.9 

Total Project Investment (TPI) $148,700,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $3,400,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.184 Other Raw Matl. Costs $500,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity -$600,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $10,000,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,400,000 

Current Yield (Actual/Theoretical) 49% Average Income Tax $6,000,000 

Average Return on Investment $14,000,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 41153 

Overall Plant Efficiency - HHV 57.0% Electricity Produced Onsite (KW) -42624 

Overall Plant Efficiency - LHV 51.6% Electricity Purchased from Grid (KW) -1471 

Plant Electricity Use (KWh/kg H2) 5.92 

Plant Steam Use (kg steam/kg H2) 19.5 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Goal Case - dry feedstock to 20% moisture content (lower gasifier temp) (Case DD)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.26 $8.77 ($/GJ H2, HHV basis) 

$10.47 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 56.8 68.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 73.5 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $19,200,000 Feedstock 40.8

 Gasification, Tar Reforming, & Quench $18,700,000 Natural Gas 4.8

 Compression & Sulfur Removal $16,000,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $22,200,000 Olivine 6.8

 Hydrogen Compression $2,700,000 Other Raw Materials 0.8

 Steam System and Power Generation $14,400,000 Waste Disposal 1.3

 Cooling Water and Other Utilities $3,600,000 Electricity 5.8 

Total Installed Equipment Cost $96,800,000 Fixed Costs 17.4 

Capital Depreciation 12.9

 Indirect Costs 49,400,000 Average Income Tax 10.1

 (% of TPI) 33.8% Average Return on Investment 24.2 

Total Project Investment (TPI) $146,100,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $2,700,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.184 Other Raw Matl. Costs $500,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $3,300,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $9,900,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,300,000 

Current Yield (Actual/Theoretical) 47% Average Income Tax $5,800,000 

Average Return on Investment $13,800,000 

Gasifier Efficiency - HHV 70.14% 

Gasifier Efficiency - LHV 69.83% Total Plant Electricity Usage (KW) 40276 

Overall Plant Efficiency - HHV 52.9% Electricity Produced Onsite (KW) -32052 

Overall Plant Efficiency - LHV 47.4% Electricity Purchased from Grid (KW) 8224 

Plant Electricity Use (KWh/kg H2) 5.96 

Plant Steam Use (kg steam/kg H2) 19.2 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Goal Case - No hydrogen recycle to the PSA (Case EE)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.21 $8.44 ($/GJ H2, HHV basis) 

$10.08 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 61.9 74.8 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 80.2 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 37.4

 Gasification, Tar Reforming, & Quench $17,600,000 Natural Gas 9.1

 Compression & Sulfur Removal $16,100,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $23,100,000 Olivine 6.2

 Hydrogen Compression $2,900,000 Other Raw Materials 0.7

 Steam System and Power Generation $14,300,000 Waste Disposal 1.1

 Cooling Water and Other Utilities $3,500,000 Electricity 6.7 

Total Installed Equipment Cost $96,400,000 Fixed Costs 15.9 

Capital Depreciation 11.8

 Indirect Costs 49,100,000 Average Income Tax 9.3

 (% of TPI) 33.7% Average Return on Investment 22.1 

Total Project Investment (TPI) $145,500,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $5,600,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.184 Other Raw Matl. Costs $400,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $4,100,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $9,900,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,300,000 

Current Yield (Actual/Theoretical) 52% Average Income Tax $5,700,000 

Average Return on Investment $13,700,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 40564 

Overall Plant Efficiency - HHV 54.7% Electricity Produced Onsite (KW) -30241 

Overall Plant Efficiency - LHV 49.1% Electricity Purchased from Grid (KW) 10322 

Plant Electricity Use (KWh/kg H2) 5.51 

Plant Steam Use (kg steam/kg H2) 18.4 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity on Goal Case - Increase PSA cost (Case FF)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$

Minimum Hydrogen Selling Price ($/kg)$1.28 $8.94 ($/GJ H2, HHV basis) 

$10.67 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 39.7

 Gasification, Tar Reforming, & Quench $17,600,000 Natural Gas 5.9

 Compression & Sulfur Removal $16,100,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $30,200,000 Olivine 6.6

 Hydrogen Compression $2,800,000 Other Raw Materials 0.7

 Steam System and Power Generation $14,200,000 Waste Disposal 1.2

 Cooling Water and Other Utilities $3,400,000 Electricity 7.1 

Total Installed Equipment Cost $103,200,000 Fixed Costs 17.6 

Capital Depreciation 13.4

 Indirect Costs 52,600,000 Average Income Tax 10.4 

(% of TPI) 33.8% Average Return on Investment 25.0 

Total Project Investment (TPI) $155,800,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $3,400,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.183 Other Raw Matl. Costs $400,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $4,100,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $10,300,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,800,000 

Current Yield (Actual/Theoretical) 49% Average Income Tax $6,100,000 

Average Return on Investment $14,600,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 40259 

Overall Plant Efficiency - HHV 53.3% Electricity Produced Onsite (KW) -29974 

Overall Plant Efficiency - LHV 47.8% Electricity Purchased from Grid (KW) 10284 

Plant Electricity Use (KWh/kg H2) 5.79 

Plant Steam Use (kg steam/kg H2) 19.5 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Goal Case - Increase in natural gas price (Case GG)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.26 $8.78 ($/GJ H2, HHV basis) 

$10.49 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 39.7

 Gasification, Tar Reforming, & Quench $17,600,000 Natural Gas 7.8

 Compression & Sulfur Removal $16,100,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $22,600,000 Olivine 6.6

 Hydrogen Compression $2,800,000 Other Raw Materials 0.7

 Steam System and Power Generation $14,200,000 Waste Disposal 1.2

 Cooling Water and Other Utilities $3,400,000 Electricity 7.1 

Total Installed Equipment Cost $95,600,000 Fixed Costs 16.8 

Capital Depreciation 12.3

 Indirect Costs 48,800,000 Average Income Tax 9.8

 (% of TPI) 33.8% Average Return on Investment 23.4 

Total Project Investment (TPI) $144,400,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $4,500,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.184 Other Raw Matl. Costs $400,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $4,100,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $9,800,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,200,000 

Current Yield (Actual/Theoretical) 49% Average Income Tax $5,700,000 

Average Return on Investment $13,700,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 40259 

Overall Plant Efficiency - HHV 53.3% Electricity Produced Onsite (KW) -29974 

Overall Plant Efficiency - LHV 47.8% Electricity Purchased from Grid (KW) 10284 

Plant Electricity Use (KWh/kg H2) 5.79 

Plant Steam Use (kg steam/kg H2) 19.5 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensivity on Goal Case - Increase in Tar Reformer/Catalyst Regeneration Cost (Case HH)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg)$1.27 $8.88 ($/GJ H2, HHV basis) 

$10.60 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 58.4 70.6 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 75.7 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 39.7

 Gasification, Tar Reforming, & Quench $17,600,000 Natural Gas 5.9

 Compression & Sulfur Removal $16,100,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $28,600,000 Olivine 6.6

 Hydrogen Compression $2,800,000 Other Raw Materials 0.7

 Steam System and Power Generation $14,200,000 Waste Disposal 1.2

 Cooling Water and Other Utilities $3,400,000 Electricity 7.1 

Total Installed Equipment Cost $101,600,000 Fixed Costs 17.4 

Capital Depreciation 13.2

 Indirect Costs 51,900,000 Average Income Tax 10.3

 (% of TPI) 33.8% Average Return on Investment 24.6 

Total Project Investment (TPI) $153,500,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $3,400,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.183 Other Raw Matl. Costs $400,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $4,100,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $10,200,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,700,000 

Current Yield (Actual/Theoretical) 49% Average Income Tax $6,000,000 

Average Return on Investment $14,400,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 40259 

Overall Plant Efficiency - HHV 53.3% Electricity Produced Onsite (KW) -29974 

Overall Plant Efficiency - LHV 47.8% Electricity Purchased from Grid (KW) 10284 

Plant Electricity Use (KWh/kg H2) 5.79 

Plant Steam Use (kg steam/kg H2) 19.5 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Goal Case - Increase steam to shift (Case II)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.28 $8.90 ($/GJ H2, HHV basis) 

$10.63 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 59.5 71.9 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 77.1 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs Operating Costs (cents/kg hydrogen)

 Feed Handling & Drying $18,900,000 Feedstock 38.9

 Gasification, Tar Reforming, & Quench $17,400,000 Natural Gas 7.6

 Compression & Sulfur Removal $16,100,000 Catalysts 0.6

 Tar Reforming Catalyst Regeneration, Shift, and PSA $23,500,000 Olivine 6.5

 Hydrogen Compression $2,800,000 Other Raw Materials 0.6

 Steam System and Power Generation $14,500,000 Waste Disposal 1.2

 Cooling Water and Other Utilities $3,100,000 Electricity 10.6 

Total Installed Equipment Cost $96,300,000 Fixed Costs 16.6 

Capital Depreciation 12.3

 Indirect Costs 49,100,000 Average Income Tax 9.7

 (% of TPI) 33.7% Average Return on Investment 23.1 

Total Project Investment (TPI) $145,500,000 Operating Costs ($/yr) 

Feedstock $23,200,000 

Natural Gas $4,500,000 

Loan Rate N/A Catalysts $400,000 

Term (years) N/A Olivine $3,800,000 

Capital Charge Factor 0.184 Other Raw Matl. Costs $400,000 

Waste Disposal $700,000 

Maximum Yields (100% of Theoretical) based on composition Electricity $6,300,000

 Theoretical Hydrogen Production (MM kg/yr) 119.7 Fixed Costs $9,900,000

 Theoretical Yield (kg/dry ton) 155.0 Capital Depreciation $7,300,000 

Current Yield (Actual/Theoretical) 50% Average Income Tax $5,800,000 

Average Return on Investment $13,700,000 

Gasifier Efficiency - HHV 72.14% 

Gasifier Efficiency - LHV 71.78% Total Plant Electricity Usage (KW) 40065 

Overall Plant Efficiency - HHV 52.0% Electricity Produced Onsite (KW) -24271 

Overall Plant Efficiency - LHV 46.5% Electricity Purchased from Grid (KW) 15793 

Plant Electricity Use (KWh/kg H2) 5.66 

Plant Steam Use (kg steam/kg H2) 23.6 
Hydrogen Production Process Engineering Analysis 

Design Report: Sensitivity for Goal Case- Decrease steam to shift (Case JJ)

2000 Dry Metric Tonnes Biomass per Day

BCL Gasifier, Tar Reformer, Sulfur Removal, HTS & LTS, PSA, Steam-Power Cycle

All Values in 2002$ 

Minimum Hydrogen Selling Price ($/kg) $1.22 $8.55 ($/GJ H2, HHV basis) 

$10.21 ($/GJ H2, LHV basis) 

Hydrogen Production at operating capacity (MM kg / year) 56.9 68.7 (Million SCF / day) 

Hydrogen Yield (kg / Dry US Ton Feedstock) 73.6 2,116 (dry tons / day) 

Delivered Feedstock Cost $/Dry US Ton $30 at operating capacity 

Internal Rate of Return (After-Tax) 10% 

Equity Percent of Total Investment 100% 

Capital Costs 

Feed Handling & Drying $18,900,000 

Gasification, Tar Reforming, & Quench $17,600,000 

Compression & Sulfur Removal $16,100,000 

Tar Reforming Catalyst Regeneration, Shift, and PSA $22,000,000 

Operating Costs (cents/kg hydrogen)

Feedstock 40.7

Natural Gas 3.2

Catalysts 0.6

Olivine 6.8

Other Raw Materials 0.8

Waste Disposal 1.2

Electricity 5.7 

Fixed Costs 17.2 

Capital Depreciation 12.7

Average Income Tax 9.9

Average Return on Investment 23.6 

Operating Costs ($/yr) 

Hydrogen Compression 

Steam System and Power Generation 

Cooling Water and Other Utilities 

Total Installed Equipment Cost 

Indirect Costs

(% of TPI)

Total Project Investment (TPI)

Loan Rate

Term (years)

Capital Charge Factor

$2,700,000 

$13,900,000 

$3,600,000 

$94,800,000 

48,300,000 

33.8% 

$143,100,000 

N/A 

N/A 

0.183 

Feedstock 

Natural Gas 

Catalysts 

Olivine 

Other Raw Matl. Costs 

Waste Disposal 

Electricity 

Fixed Costs 

Capital Depreciation 

Average Income Tax 

Average Return on Investment 

Total Plant Electricity Usage (KW) 

Electricity Produced Onsite (KW) 

Electricity Purchased from Grid (KW) 

Plant Electricity Use (KWh/kg H2) 

Plant Steam Use (kg steam/kg H2) 

$23,200,000 

$1,800,000 

$400,000 

$3,800,000 

$400,000 

$700,000 

$3,300,000

$9,800,000

$7,200,000 

$5,600,000 

$13,400,000 

40210 

-32059 

8151 

5.95 

17.8 

Maximum Yields (100% of Theoretical) based on composition 

Theoretical Hydrogen Production (MM kg/yr) 

Theoretical Yield (kg/dry ton) 

Current Yield (Actual/Theoretical) 

Gasifier Efficiency - HHV 

Gasifier Efficiency - LHV 

Overall Plant Efficiency - HHV 

Overall Plant Efficiency - LHV 

119.7 

155.0 

48% 

72.14% 

71.78% 

53.7%

48.1%
Appendix C: Current Design Process Flow Diagrams 
















Appendix D: Goal Design Process Flow Diagrams 















Appendix E: Graphical Correlations for Gas Components and Char 
BCL H2 Correlation 

y = 1.8930E-05x2 - 2.6448E-02x + 1.7996E+01 

R2

 = 9.1780E-01 

0 

5 

10 

15 

20 

25 

30 

35 

40 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

mol% of dry gas 

BCL CO Correlation 

y = 2.8792E-05x2 - 1.0290E-01x + 1.3346E+02 

R2

 = 3.9738E-01 

0 

10 

20 

30 

40 

50 

60 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

mol% of dry gas 

Temp (F) 
BCL CO2 Correlation 

y = -1.4927E-05x2 + 3.7889E-02x - 9.5251E+00 

R2

 = 4.1740E-01 

0 

2 

4 

6 

8 

10 

12 

14 

16 

18 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

mol% of dry gas 

BCL CH4 Correlation 

y = -1.6167E-05x2 + 4.4179E-02x - 1.3820E+01 

R2

 = 6.9332E-01 

0 

2 

4 

6 

8 

10 

12 

14 

16 

18 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

mol% of dry gas 
BCL C2H2 Correlation 

(Note regressed data with values greater than zero) 

y = -1.5610E-06x2 + 5.4499E-03x - 4.3114E+00 

R2 = 7.1662E-01 

0.00 

0.10 

0.20 

0.30 

0.40 

0.50 

0.60 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

mol% of dry gas 

BCL C2H4 Correlation 

y = -1.9868E-05x2 + 5.8435E-02x - 3.8258E+01 

R2

 = 8.7617E-01 

0 

1 

2 

3 

4 

5 

6 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

mol% of dry gas 
BCL C2H6 Correlation 

y = 3.0640E-06x2 - 1.1667E-02x + 1.1114E+01 

R2

 = 8.5202E-01 

0 

0.2 

0.4 

0.6 

0.8 

1 

1.2 

1.4 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

mol% of dry gas 

BCL Char Correlation 

y = -3.1178E-08x2 - 3.0212E-04x + 7.5503E-01 

R2

 = 6.5481E-01 

0 

0.05 

0.1 

0.15 

0.2 

0.25 

0.3 

0.35 

0.4 

0.45 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

lb/lb of dry wood 
BCL Tar Correlation 

y = 6.3527E-22x2 - 1.9759E-05x + 4.5494E-02 

R2

 = 1.0000E+00 

0.000 

0.005 

0.010 

0.015 

0.020 

0.025 

0 200 400 600 800 1,000 1,200 1,400 1,600 1,800 2,000 

Temp (F) 

lb/lb of dry wood 
Appendix F: Flow Charts for Gasifier Elemental Balances 
BCL model Fortran 

 performs 

balances in the following order: 

1. Carbon 

2. Oxygen 

3. Sulfur BCL model - Carbon balance 

4. Hydrogen 

where TCARB = carbon in 

the wood 

gaseous components (CO + 

CH4 + .....) 

Is CCARB < 0? 

no 

yes 

Calc char carbon (CARB) 

CCARB = TCARB – GCARB 

Calc gaseous carbon (GCARB) 

GCARB = sum of C in all 

Carbon balances 

Gaseous carbon, GCARB, 

exceeds carbon available 

1. Set CCARB = 0 

2. Adjust molar flow; multiply by 

CFACT = TCARB/GCARB 

Carbon balances 
BCL model - Oxygen balance 

yes 

) 

no 

Calc oxygen in the char (COXY) 

Based on Real(6) - % O in wood that 

is retained in the char 

Is OREQD 

>TOXY? 

Calc oxygen required: 

OREQD = O in gaseous components (CO + 

CO2) + O in char (COXY

Calc oxygen available: 

OAVAIL = O in 

conventional stream as 

O2 + O in wood 

(TOXY) + H2OTOT (O 

in moisture in wood + O 

in steam to gasifier) Calculate excess O: 

OXY = TOXY – OREQD 

Add excess O to char 

COXY = COXY + OXY 

Oxygen balances 

Decompose steam to elements and 

reduce amount of steam: 

STMDEC = OREQD – TOXY 

H2OTOT = H2OTOT - STMDEC 

OXY = 0 

< 0? 

yes 

no 

Write error: need 

Is H2OTOT 

more steam 

Oxygen balances 
BCL model - Sulfur balance 

Assume all gaseous sulfur is present as 

H2S AND all solid sulfur appears in the 

char 

Calc H2S and sulfur in char (CSULF) 

using variable Real(4) - % wood sulfur 

retained in the char 

TSULF = sulfur in the wood 

H2SMOL = TSULF *(1-Real(4)/100) 

CSULF = TSULF*(Real(4)/100) 

Sulfur balances 
BCL model 

Hydrogen balance 

Using Real(3) - % wood nitrogen 

retained in char calc: 

1. CNIT – nitrogen in the char 

(used later in char output) 

2. AMMOL – remaining nitrogen 

goes to NH3 

components (CH4 + C2H4 + .....) 

HREQD = GASHYD + CHRHYD 

hydrogen in the char 

Is HREQD > 

HAVAIL? 

yield exceeds 

should be increased 

yes (HAVAIL) 

HAVAIL = H in wood 

O balance 

no 

Is HREQD > 

THYD (H in 

wood)? 

yes 

) 

XSHYD = (2*STMDEC) – 

H2STM 

where: H2STM is excess H 

that is not in the wood; no 

Calc gaseous hydrogen (GASHYD) 

GASHYD = sum of H in all gaseous 

Calc hydrogen required (HREQD) 

Using Real(5) – H/C 

mass ratio in char calc: 

CHRHYD – amount of 

Write error: hydrogen 

hydrogen available; 

stm to wood ratio 

Calc hydrogen available 

(THYD) + H from any 

steam decomposition in 

Calc excess H that does not 

come from steam 

decomposition (XSHYD

H2STM = HREQD - THYD 

STMDEC is the steam 

Calc excess H (HYD) 

HYD = (HAVAIL - HREQ) 

Convert excess H to char 

CHRHYD = HYD + CHRHYD 

Hydrogen balances 

decomposed in O balance 
Appendix G: Equipment Design Parameters and Cost References 
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD-P700-A101-2 

C-101 Hopper Feeder CONVEYOR VIBRATING-FEEDER Included in overall cost for feed handling & drying taken from several literature sources C-102 Screener Feeder Conveyor CONVEYOR BELT Included in overall cost for feed handling & drying taken from several literature sources 

C-103 Radial Stacker Conveyor CONVEYOR BELT Included in overall cost for feed handling & drying taken from several literature sources C-104 Dryer Feed Screw Conveyor CONVEYOR SCREW Included in overall cost for feed handling & drying taken from several literature sources C-105 Gasifier Feed Screw Conveyor CONVEYOR SCREW Included in overall cost for feed handling & drying taken from several literature sources H-101 Flue Gas Cooler / Steam Generator #3 HEATX SHELL-TUBE duty = 1.37 MMBtu/hr; LMTD = 1,220 F; U = 150 Btu/hr-ft^2-F; area = 7 ft^2; fixed TS K-101 Flue Gas Blower FAN CENTRIFUGAL Included in overall cost for feed handling & drying taken from several literature sources M-101 Hydraulic Truck Dump with Scale SCALE TRUCK-SCALE Included in overall cost for feed handling & drying taken from several literature sources 

M-102 Hammermill SIZE-REDUCTION Included in overall cost for feed handling & drying taken from several literature sources M-103 Front End Loaders VEHICLE LOADER Included in overall cost for feed handling & drying taken from several literature sources 

M-104 Rotary Biomass Dryer DRYER ROTARY-DRUM Included in overall cost for feed handling & drying taken from several literature sources 

S-101 Magnetic Head Pulley SEPARATOR MAGNET Included in overall cost for feed handling & drying taken from several literature sources S-102 Screener SEPARATOR SCREEN Included in overall cost for feed handling & drying taken from several literature sources S-103 Dryer Air Cyclone SEPARATOR GAS CYCLONE Included in overall cost for feed handling & drying taken from several literature sources S-104 Dryer Air Baghouse Filter SEPARATOR FABRIC-FILTER Included in overall cost for feed handling & drying taken from several literature sources 

T-101 Dump Hopper TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literature sources T-102 Hammermill Surge Bin TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literature sources 

T-103 Dryer Feed Bin TANK LIVE-BTM-BIN Included in overall cost for feed handling & drying taken from several literature sources 

T-104 Dried Biomass Hopper TANK VERTICAL-VESSEL Included in overall cost for feed handling & drying taken from several literature sources 

PFD-P700-A201-2 

C-201 Sand/ash Conditioner/Conveyor CONVEYOR SCREW Included in overall cost for gasification & gas clean up taken from several literature sources H-201 Post-tar Reformer Cooler / Steam Generator #1 HEATX SHELL-TUBE duty = 47.9 MMBtu/hr; LMTD = 457; area = 698 sq ft; U = 150 Btu/hr-ft^2-F; fixed TS 

H-202 Post-tar Reformer Cooler / BFW Preheater #2 HEATX SHELL-TUBE duty = 79.4 MMBTU; LMTD = 133 F; U = 150 Btu/hr-ft^2-F; area = 5,946 ft^2; fixed TS K-201 Combustion Air Blower FAN CENTRIFUGAL Included in overall cost for gasification & gas clean up taken from several literature sources M-201 Sand/ash Cooler MISCELLANEOUS MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources R-201 Indirectly-heated Biomass Gasifier REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literature sources R-202 Char Combustor REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literature sources R-203 Tar Reformer REACTOR VERTICAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literature sources S-201 Primary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several literature sources S-202 Secondary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several literature sources S-203 Primary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several literature sources 

S-204 Secondary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up taken from several literature sources 

S-205 Electrostatic Precipitator SEPARATOR MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources T-201 Sand/ash Bin TANK FLAT-BTM-STORAGE Included in overall cost for gasification & gas clean up taken from several literature sources 

PFD-P700-A301-3 

H-301 Quench Water Recirculation Cooler HEATX SHELL-TUBE Included in overall cost for gasification & gas clean up taken from several literature sources 

H-302 Syngas Compressor Intercoolers HEATX AIR-COOLED EXCHANGER Cost of intercoolers included in cost for syngas compressor, K-301 

H-303 Water-cooled Aftercooler HEATX SHELL-TUBE duty = 2.9 MMBtu/hr; LMTD = 25F; U = 150 Btu/hr-ft^2-F; surface area = 794 ft^2; fixed TS 

H-304 LO-CAT Preheater HEATX SHELL-TUBE duty = 0.8 MMBtu/hr;LMTD = 87 F; U = 90 Btu/hr-ft^2-F; surface area = 98 ft^2; fixed TS H-305 LO-CAT Absorbent Solution Cooler HEATX SHELL-TUBE Included in LO-CAT system cost H-306 ZnO Bed Preheater HEATX SHELL-TUBE duty = 47 MMBtu/hr duty; LMTD = 102 F; U = 90 Btu/hr-ft^2-F; area = 5,137 ft^2; fixed TS K-301 Syngas Compressor COMPRESSOR CENTRIFUGAL gas flow rate = 70,000 CFM; 6 impellers; design outlet pressure = 465 psi; 30,000 HP; intercoolers, aftercooler, & K.O.s incluK-302 LO-CAT Feed Air Blower FAN CENTRIFUGAL Included in LO-CAT system cost K-303 Reformer Flue Gas Blower FAN CENTRIFUGAL gas flow rate (actual) = 148,464 CFM; 327 HP M-301 Syngas Quench Chamber MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources M-302 Syngas Venturi Scrubber MISCELLANEOUS Included in overall cost for gasification & gas clean up taken from several literature sources M-303 LO-CAT Venturi Precontactor MISCELLANEOUS Included in LO-CAT system cost M-304 LO-CAT Liquid-filled Absorber COLUMN ABSORBER Included in LO-CAT system cost P-301 Sludge Pump PUMP CENTRIFUGAL 1.4 GPM; 0.053 brake HP; design pressure = 60 psia P-302 Quench Water Recirculation Pump PUMP CENTRIFUGAL Included in the cost of the gasification & gas clean up system P-303 LO-CAT Absorbent Solution Circulating Pump PUMP CENTRIFUGAL Included in LO-CAT system cost R-301 LO-CAT Oxidizer Vessel REACTOR VERTICAL-VESSEL Included in LO-CAT system cost R-302 ZnO Sulfur Removal Beds REACTOR VERTICAL-VESSEL 6 ft diameter; 13 ft height; 427 cub ft volume; 490 psia design pressure; 757 F design temperature 
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

S-301 Pre-compressor Knock-out SEPARATOR KNOCK-OUT DRUM 18 ft diameter; 36 ft height; design pres = 40 psia; design temp = 197 F S-302 Syngas Compressor Interstage Knock-outs SEPARATOR KNOCK-OUT DRUM Cost of intercoolers K.O.s included in cost for syngas compressor, K-301 S-303 Post-compressor Knock-out SEPARATOR KNOCK-OUT DRUM 7 ft. diameter; 14 ft height; design pres = 506 psia; design temp = 160 F T-301 Sludge Settling Tank SEPARATOR CLARIFIER 3 ft diameter; 7 ft height; 431 gal volume; T-302 Quench Water Recirculation Tank TANK HORIZONTAL-VESSEL Included in overall cost for gasification & gas clean up taken from several literature sources 

PFD-P700-A401-3 

H-401 Reformer Feed Preheater HEATX SHELL-TUBE duty = 47.6 MMBtu/hr; LMTD = 491 F; U = 90 Btu/hr-ft^2-F; area = 1,078 ft^2; fixed TS H-402 Reformed Syngas Cooler/Steam Generator #2 HEATX SHELL-TUBE duty = 155 MMBtu/hr; LMTD = 733 F; U = 150 Btu/hr-ft^2-F; area = 1,410 ft^2; fixed tube sheet H-403 Reformed Syngas Cooler/Steam Superheater #1 HEATX SHELL-TUBE duty = 14 MMBtu/hr; U = 150 Btu/hr-ft^2-F; area = 983 ft^2; LMTD = 95 F H-404 Reformer Flue Gas Cooler/Steam Superheater #2 HEATX SHELL-TUBE duty = 94 MMBtu/hr; LMTD = 217 F; U = 150 Btu/hr-ft^2-F; area = 2,900 ft^2; fixed TS H-405 LT Shift Precooler/BFW Preheater #1 HEATX SHELL-TUBE duty = 54 MMBtu/hr; LMTD = 249 F; U = 100 Btu/hr-ft^2-F; area = 2,190 ft^2; fixed TS H-406 LT shift Precooler/Deaerator Water Preheater #1 HEATX SHELL-TUBE duty = 20 MMBtu/hr; LMTD = 244 F; U = 100 Btu/hr-ft^2-F; area = 823 ft^2; fixed TS H-407 PSA Precooler / Deaerator Water Preheater #2 HEATX SHELL-TUBE duty = 21 MMBtu/hr; LMTD = 251 F; U = 100 Btu/hr-ft^2-F; area = 858 ft^2; fixed TS H-408 PSA Air-cooled Precooler HEATX AIR-COOLED EXCHANGER duty = 149 MMBtu/hr; LMTD = 103 F; U = 90 Btu/hr-ft^2-F; area = 16,117 ft^2; air cooler H-409 PSA Water-cooled Precooler HEATX SHELL-TUBE duty = 8 MMBtu/hr; LMTD = 25 F; U = 150 Btu/hr-ft^2-F; surface area = 2,274 ft^2 K-401 Reformer Combustion Air Blower FAN CENTRIFUGAL gas flow rate (actual) = 70133 CFM; outlet pressure = 9.88 inches H2O R-401 Steam Reformer REACTOR VERTICAL-VESSEL heat duty = 159 MMBtu/hr R-402 High Temperature Shift Reactor REACTOR VERTICAL-VESSEL GHSV = 3,000/hr; H/D = 2; 12 ft diameter; 24 ft height; 400 psia op press; 807 F op temp R-403 Low Temperature Shift Reactor REACTOR VERTICAL-VESSEL GHSV = 4,000; H/D = 2; 11 ft diameter; 22 ft height; 390 psia op press; 453 F op temp S-401 Pre-PSA Knock-out #1 SEPARATOR KNOCK-OUT DRUM H/D = 2; 12 ft diameter; 23 ft height; operating pressure = 380 psi; operating temperature = 334 F S-402 Pre-PSA Knock-out #2 SEPARATOR KNOCK-OUT DRUM H/D = 2; 9 ft diameter; 17 ft height; operating pressure = 370 psi; operating temperature = 110 F S-403 Pressure Swing Adsorption Unit MISCELLANEOUS PACKAGE several beds; cost scaled from value of $0.168/SCFD of H2 

PFD-P700-A501 

H-501A Hydrogen Compressor Intercooler HEATX AIR-COOLED EXCHANGER duty = 4 MMBtu/hr; LMTD = 61 F; U = 90 Btu/hr-ft^2-F; area = 740 ft^2; air cooler H-501B Hydrogen Compressor Air-cooled Aftercooler HEATX AIR-COOLED EXCHANGER duty = 6 MMBtu/hr; LMTD = 77 F; U = 90 Btu/hr-ft^2-F; area = 864 sq ft.; air cooler H-502 Hydrogen Compressor Water-cooler Aftercooler HEATX SHELL-TUBE duty = 1.5 MMBtu/hr; LMTD = 25 F; U = 150 Btu/hr-ft^2-F; area = 396 ft^2; fixed TS K-501 Hydrogen Compressor COMPRESSOR RECIPROCATING gas flow rate = 2,028 actual CFM; outlet pressure = 1,020 psi S-501 Pre-hydrogen Compressor Knock-out SEPARATOR KNOCK-OUT DRUM H/D = 2; 3 ft diam; 7 ft height; operating pressure = 360 psia; operating temperature = 109 F S-502 Hydrogen Compressor 1st Interstage Knock-out SEPARATOR KNOCK-OUT DRUM included in the price of the hydrogen compressor (K-501) S-503 Post-hydrogen Compressor Knock-out SEPARATOR KNOCK-OUT DRUM H/D = 2; 3 ft diameter; 5 ft height;; operating pressure = 1,015 psi; operating temperature = 110 F 

PFD-P700-A601-3 

H-601 Steam Turbine Condenser HEATX SHELL-TUBE Included in the cost of the steam trubine/generator (M-602); condenser steam flow rate = 342,283 lb/hr H-602 Blowdown Cooler / Deaerator Water Preheater HEATX SHELL-TUBE duty = 3 MMBtu/hr; LMTD = 236 F; U = 600 Btu/hr-ft^2-F; area = 20 ft^2; pre-engineered U-tube H-603 Blowdown Water-cooled Cooler HEATX SHELL-TUBE duty = 0.6 MMBtu/hr; LMTD = 47 F; U = 225 Btu/hr-ft^2-F; area = 60 ft^2; fixed TS M-601 Hot Process Water Softener System MISCELLANEOUS PACKAGE scaled cost to 700 gpm flow, 24" dia softener. Includes filters, chemical feeders, piping, valves 

M-602 Extraction Steam Turbine/Generator GENERATOR STEAM-TURBINE 25.6 MW generated; 34,308 HP P-601 Collection Pump PUMP CENTRIFUGAL 513 GPM; 4 brake HP; outlet pressure = 25 psia P-602 Condensate Pump PUMP CENTRIFUGAL 190 GPM; 4 brake HP; outlet pressure = 25 psia 

P-603 Deaerator Feed Pump PUMP CENTRIFUGAL 702 GPM; 14 brake HP; outlet pressure = 40 psia P-604 Boiler Feed Water Pump PUMP CENTRIFUGAL 730 GPM; 759 brake HP; outlet pressure = 1,345 psia S-601 Blowdown Flash Drum TANK HORIZONTAL-VESSEL H/D = 2; residence time = 5 min; 2 ft diameter; 4 ft height; op press = 1,280 psi; op temp = 575 F T-601 Condensate Collection Tank TANK HORIZONTAL-VESSEL residence time = 10 minutes; H/D = 2; 8 ft diameter; 17 ft height T-602 Condensate Surge Drum TANK HORIZONTAL-VESSEL residence time = 10 minutes; H/D = 2; 9 ft diameter; 17 ft height T-603 Deaerator TANK HORIZONTAL-VESSEL liquid flow rate = 348,266 lb/hr; 150 psig design pressure; 10 min residence time T-604 Steam Drum TANK HORIZONTAL-VESSEL 424 gal, 4.5' x 4'dia, 15 psig 

PFD-P700-A701-2 

K-701 Plant Air Compressor COMPRESSOR RECIPROCATING 450 cfm, 125 psig outlet M-603 Startup Boiler MISCELLANEOUS PACKAGE Assume need steam requirement equal to 1/2 of steam requirement for gasifier at full rate steam rate = 36,560 lb/hr 

M-701 Cooling Tower System COOLING-TOWER INDUCED-DRAFT approx 16,500 gpm, 140 MMBtu/hr M-702 Hydraulic Truck Dump with Scale SCALE TRUCK-SCALE Hydraulic Truck Dumper with Scale M-703 Flue Gas Stack MISCELLANEOUS MISCELLANEOUS 42 inch diameter; 250 deg F P-701 Cooling Water Pump PUMP CENTRIFUGAL 16,188 GPM; 659 brake HP; outlet pressure 75 psi P-702 Firewater Pump PUMP CENTRIFUGAL 2,500 gpm, 50 ft head P-703 Diesel Pump PUMP CENTRIFUGAL 30 gpm, 150 ft head 
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

P-704 Ammonia Pump PUMP CENTRIFUGAL 8.5 gpm, 22 ft head P-705 Hydrazine Pump PUMP CENTRIFUGAL 5 gpm, 75 ft head S-701 Instrument Air Dryer DRYER PACKAGE 400 SCFM Air Dryer, -40 F Dewpoint T-701 Plant Air Receiver TANK HORIZONTAL-VESSEL 900 gal., 200 psig T-702 Firewater Storage Tank TANK FLAT-BTM-STORAGE 600,000 gal, 4 hr res time, 51' dia x 40' high, atmospheric T-703 Diesel Storage Tank TANK FLAT-BTM-STORAGE 10,667 gal, 120 hr res time, 90% wv, 10' dia x 18.2' high, atmospheric T-704 Ammonia Storage Tank TANK HORIZONTAL-STORAGE Included in the cost of the feed handling step. T-705 Olivine Lock Hopper TANK VERTICAL-VESSEL Included in the cost of the feed handling step. T-706 MgO Lock Hopper TANK VERTICAL-VESSEL 20' x 20' Bin, Tapering to 3' x 3' at Bottom. Capacity 6,345 cf, two truck loads. T-707 Hydrazine Storage Tank TANK VERTICAL-VESSEL 260 gal, 4.9' x 3'dia., 10psig 
Appendix H: Current Design Summary of Individual Equipment Costs 
Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name Scaling Stream 

Scaling Stream 

Flow (lb/hr or 

btu/hr) 

New Stream 

Flow Size Ratio 

Original Equip 

Cost (per unit) Base Year 

Total Original Equip 

Cost (Req'd & Spare) 

in Base Year 

Scaling 

Exponent 

Scaled Cost in 

Base Year 

In

C-101 4 Hopper Feeder 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-102 2 Screener Feeder Conveyor 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-103 2 Radial Stacker Conveyor 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-104 2 Dryer Feed Screw Conveyor 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-105 2 Gasifier Feed Screw Conveyor 104 208,771 208,771 1.00 Included in feed handling & drying cost (M-104) H-101 1 Flue Gas Cooler / Steam Generator #3 PINCH 1,369,986 1,369,986 1.00 $26,143 2002 $26,143 0.6 $26,143 K-101 2 Flue Gas Blower 112 639,530 639,530 1.00 Included in feed handling & drying cost (M-104) M-101 4 Hydraulic Truck Dump with Scale 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) M-102 2 Hammermill 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) M-103 3 Front End Loaders 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) M-104 2 Rotary Biomass Dryer 101 367,437 367,437 1.00 $3,813,728 2002 $7,627,455 0.75 $7,627,450 

S-101 2 Magnetic Head Pulley 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) S-102 2 Screener 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) S-103 2 Dryer Air Cyclone 111 639,530 639,530 1.00 Included in feed handling & drying cost (M-104) S-104 2 Dryer Air Baghouse Filter 103 208,771 208,771 1.00 Included in feed handling & drying cost (M-104) T-101 4 Dump Hopper 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) T-102 1 Hammermill Surge Bin 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) T-103 2 Dryer Feed Bin 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) T-104 2 Dried Biomass Hopper 104 208,771 208,771 1.00 Included in feed handling & drying cost (M-104) A100 Subtotal $7,653,598 $7,653,593 

C-201 1 Sand/ash Conditioner/Conveyor 219 7,380 47,912,711 1.00 Included in gasification & clean up cost (R-201) H-201 1 Post-tar Reformer Cooler / Steam Generator #1 PINCH 47,912,711 79,370,881 1.00 $69,089 2002 $69,089 0.65 $69,089 

H-202 1 Post-tar Reformer Cooler / BFW Preheater #2 PINCH 79,370,881 79,370,881 1.00 $99,389 2002 $99,389 0.6 $99,389 

K-201 2 Combustion Air Blower 208 442,163 442,163 1.00 Included in gasification & clean up cost (R-201) M-201 2 Sand/ash Cooler 217 6,642 6,642 1.00 Included in gasification & clean up cost (R-201) R-201 2 Indirectly-heated Biomass Gasifier 201 5,228,880 5,228,878 1.00 $3,318,302 2002 $6,636,603 0.65 $6,636,601 

R-202 2 Char Combustor 210 5,434,490 5,434,493 1.00 Included in gasification & clean up cost (R-201) R-203 1 Tar Reformer 225 241,995 241,995 1.00 Included in gasification & clean up cost (R-201) S-201 2 Primary Gasifier Cyclone 202 5,228,880 5,228,878 1.00 Included in gasification & clean up cost (R-201) S-202 2 Secondary Gasifier Cyclone 222 246,484 246,483 1.00 Included in gasification & clean up cost (R-201) S-203 2 Primary Combustor Cyclone 210 5,434,490 5,434,493 1.00 Included in gasification & clean up cost (R-201) S-204 2 Secondary Combustor Cyclone 212 487,506 487,506 1.00 Included in gasification & clean up cost (R-201) 

S-205 2 Electrostatic Precipitator 213 480,870 480,870 1.00 Included in gasification & clean up cost (R-201) T-201 1 Sand/ash Bin 217 6,642 6,642 1.00 Included in gasification & clean up cost (R-201) A200 Subtotal $6,805,081 $6,805,079 
Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name Scaling Stream 

Scaling Stream 

Flow (lb/hr or 

btu/hr) 

New Stream 

Flow Size Ratio 

Original Equip 

Cost (per unit) Base Year 

Total Original Equip 

Cost (Req'd & Spare) 

in Base Year 

Scaling 

Exponent 

Scaled Cost in 

Base Year 

In

H-301 1 Quench Water Recirculation Cooler 301 241,995 241,995 1.00 Included in gasification & clean up cost (R-201) H-302 5 Syngas Compressor Intercoolers 301 241,995 241,995 1.00 Included in the syngas compressor cost (K-301) H-303 1 Water-cooled Aftercooler QCH303CT 2,938,799 2,940,165 1.00 $20,889 2002 $20,889 0.44 $20,893 H-304 1 LO-CAT Preheater PINCH 770,434 770,434 1.00 $4,743 2002 $4,743 0.6 $4,743 H-305 1 LO-CAT Absorbent Solution Cooler 320 179,394 179,394 1.00 Included in LO-CAT oxidizer vessel cost (R-301) H-306 1 ZnO Bed Preheater PINCH 47,209,942 47,209,942 1.00 $71,389 2002 $71,389 0.44 $71,389 

K-301 1 Syngas Compressor 315 220,009 220,009 1.00 $4,817,834 2002 $4,817,834 0.8 $4,817,834 

K-302 1 LO-CAT Feed Air Blower 322 359 359 1.00 Included in LO-CAT oxidizer vessel cost (R-301) K-303 1 Reformer Flue Gas Blower 434 534,677 534,677 1.00 $54,250 2002 $54,250 0.59 $54,250 

M-301 1 Syngas Quench Chamber 301 241,496 241,995 1.00 Included in gasification & clean up cost (R-201) M-302 1 Syngas Venturi Scrubber 301 241,496 241,995 1.00 Included in gasification & clean up cost (R-201) M-303 1 LO-CAT Venturi Precontactor 323 517 517 1.00 Included in LO-CAT oxidizer vessel cost (R-301) M-304 1 LO-CAT Liquid-filled Absorber 320 179,394 179,394 1.00 Included in LO-CAT oxidizer vessel cost (R-301) P-301 1 1 Sludge Pump 336 997 997 1.00 $3,911 2002 $7,822 0.33 $7,823 P-302 1 1 Quench Water Recirculation Pump 307 1,272,120 1,272,123 1.00 Included in gasification & clean up cost (R-201) P-303 1 1 LO-CAT Absorbent Solution Circulating Pump 301 241,496 241,995 1.00 Included in LO-CAT oxidizer vessel cost (R-301) R-301 1 LO-CAT Oxidizer Vessel 323 517 517 1.00 $1,000,000 2002 $1,000,000 0.65 $999,653 R-302 2 ZnO Sulfur Removal Beds 327 179,237 179,237 1.00 $37,003 2002 $74,006 0.56 $74,006 

S-301 1 Pre-compressor Knock-out 315 220,009 220,009 1.00 $157,277 2002 $157,277 0.6 $157,277 

S-302 4 Syngas Compressor Interstage Knock-outs 315 220,009 220,009 1.00 Included in the syngas compressor cost (K-301) S-303 1 Post-compressor Knock-out 319 179,394 179,394 1.00 $40,244 2002 $40,244 0.6 $40,244 T-301 1 Sludge Settling Tank 302 21,718 21,718 1.00 $11,677 2002 $11,677 0.6 $11,677 T-302 1 Quench Water Recirculation Tank 301 241,496 241,995 1.00 Included in gasification & clean up cost (R-201) A300 Subtotal $6,260,131 $6,259,790 

H-401 1 Reformer Feed Preheater QH401 47,628,665 47,628,665 1.00 $277,489 2002 $277,489 0.7 $277,489 

H-402 1 Reformed Syngas Cooler/Steam Generator #2 PINCH 155,010,823 155,010,823 1.00 $347,989 2002 $347,989 0.6 $347,989 

H-403 1 Reformed Syngas Cooler/Steam Superheater #1 PINCH 13,974,577 13,974,577 1.00 $92,889 2002 $92,889 0.6 $92,889 

H-404 1 Reformer Flue Gas Cooler/Steam Superheater #2 PINCH 94,212,763 94,212,763 1.00 $196,589 2002 $196,589 0.6 $196,589 

H-405 1 LT Shift Precooler/BFW Preheater #1 PINCH 54,476,359 54,476,359 1.00 $56,089 2002 $56,089 0.6 $56,089 

H-406 1 LT shift Precooler/Deaerator Water Preheater #1 PINCH 20,095,131 20,095,131 1.00 $20,989 2002 $20,989 0.6 $20,989 H-407 1 PSA Precooler / Deaerator Water Preheater #2 PINCH 21,034,730 21,034,730 1.00 $21,089 2002 $21,089 0.6 $21,089 H-408 1 PSA Air-cooled Precooler QAH408 149,281,592 149,281,592 1.00 $388,064 2002 $388,064 0.6 $388,064 

H-409 1 PSA Water-cooled Precooler QCH409CT 8,414,338 8,414,338 1.00 $35,689 2002 $35,689 0.44 $35,689 K-401 1 Reformer Combustion Air Blower 430 304,578 304,578 1.00 $35,020 2002 $35,020 0.59 $35,020 R-401 1 Steam Reformer QR401 158,705,747 158,705,747 1.00 $4,965,833 2002 $4,965,833 0.7 $4,965,833 

R-402 1 High Temperature Shift Reactor 404 354,424 354,424 1.00 $465,907 2002 $465,907 0.56 $465,907 

R-403 1 Low Temperature Shift Reactor 407 354,424 354,424 1.00 $323,464 2002 $323,464 0.56 $323,464 

S-401 1 Pre-PSA Knock-out #1 413 354,424 354,424 1.00 $129,979 2002 $129,979 0.6 $129,979 

S-402 1 Pre-PSA Knock-out #2 419 242,691 242,691 1.00 $55,291 2002 $55,291 0.6 $55,291 

S-403 1 Pressure Swing Adsorption Unit 424 14,260 14,260 1.00 $4,855,471 2002 $4,855,471 0.6 $4,855,482 

A400 Subtotal $12,267,841 $12,267,853 
Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name Scaling Stream 

Scaling Stream 

Flow (lb/hr or 

btu/hr) 

New Stream 

Flow Size Ratio 

Original Equip 

Cost (per unit) Base Year 

Total Original Equip 

Cost (Req'd & Spare) 

in Base Year 

Scaling 

Exponent 

Scaled Cost in 

Base Year 

In

H-501A 1 Hydrogen Compressor Intercooler QAK501A 4,042,813 4,042,813 1.00 $53,601 2002 $53,601 0.6 $53,601 

H-501B 1 Hydrogen Compressor Air-cooled Aftercooler QAK501B 5,984,714 5,984,714 1.00 $56,901 2002 $56,901 0.6 $56,901 

H-502 1 Hydrogen Compressor Water-cooler Aftercooler QCH502CT 1,465,277 1,465,278 1.00 $18,909 2002 $18,909 0.44 $18,909 K-501 1 Hydrogen Compressor 501 14,260 14,260 1.00 $914,235 2002 $914,235 0.8 $914,238 

S-501 1 Pre-hydrogen Compressor Knock-out 501 14,260 14,260 1.00 $13,377 2002 $13,377 0.6 $13,377 S-502 1 Hydrogen Compressor 1st Interstage Knock-out 501 14,260 14,260 1.00 Included in the hydrogen compressor cost (K-501) S-503 1 Post-hydrogen Compressor Knock-out 505 14,260 14,260 1.00 $13,977 2002 $13,977 0.6 $13,977 A500 Subtotal $1,071,000 $1,071,003 

H-601 1 Steam Turbine Condenser 614 93,974 93,974 1.00 Included in the extraction steam trubine/generator cost (M-602) H-602 1 Blowdown Cooler / Deaerator Water Preheater PINCH 2,877,029 2,877,029 1.00 $3,043 2002 $3,043 0.6 $3,043 H-603 1 Blowdown Water-cooled Cooler QCH603CT 626,343 626,343 1.00 $16,143 2002 $16,143 0.44 $16,143 M-601 1 Hot Process Water Softener System 631 349,266 349,266 1.00 $1,031,023 1999 $1,031,023 0.82 $1,031,023 M-602 1 Extraction Steam Turbine/Generator 607 342,283 342,283 1.00 $4,045,870 2002 $4,045,870 0.71 $4,045,870 M-603 1 Startup Boiler 200 36,560 36,560 1.00 $198,351 2002 $198,351 0.6 $198,351 

P-601 1 1 Collection Pump 625 255,292 255,292 1.00 $7,015 2002 $14,030 0.33 $14,030 P-602 1 1 Condensate Pump 616 93,974 93,974 1.00 $5,437 2002 $10,874 0.33 $10,874 

P-603 1 1 Deaerator Feed Pump 628 349,266 349,266 1.00 $8,679 2002 $17,358 0.33 $17,358 P-604 1 1 Boiler Feed Water Pump 639 349,268 349,268 1.00 $95,660 2002 $191,320 0.33 $191,320 

T-601 1 Condensate Collection Tank 627 349,266 349,266 1.00 $24,493 2002 $24,493 0.6 $24,493 T-602 1 Condensate Surge Drum 638 349,268 349,268 1.00 $28,572 2002 $28,572 0.6 $28,572 T-603 1 Deaerator 633 349,266 349,266 1.00 $130,721 2002 $130,721 0.72 $130,721 

T-604 1 Steam Drum 644 349,268 349,268 1.00 $9,200 1997 $9,200 0.72 $9,200 S-601 1 Blowdown Flash Drum 604 6,985 6,985 1.00 $14,977 2002 $14,977 0.6 $14,977 A600 Subtotal $5,735,975 $5,735,976 

K-701 2 1 Plant Air Compressor 101 367,437 367,437 1.00 $32,376 2002 $97,129 0.34 $97,129 

M-701 1 Cooling Tower System QCTOTAL 139,850,763 139,850,763 1.00 $267,316 2002 $267,316 0.78 $267,316 

M-702 1 Hydraulic Truck Dump with Scale 101 367,437 367,437 1.00 $80,000 1998 $80,000 0.6 $80,000 M-703 1 Flue Gas Stack 112 1,174,206 639,530 1.00 $51,581 2002 $51,581 1 $51,581 

434 534,677 The stack flow is the sum of two flow streams. 

P-701 1 1 Cooling Water Pump 715 6,088,320 6,113,668 1.00 $158,540 2002 $317,080 0.33 $317,515 P-702 1 1 Firewater Pump 101 367,437 367,437 1.00 $18,400 1997 $36,800 0.79 $36,800 P-703 1 1 Diesel Pump 101 367,437 367,437 1.00 $6,100 1997 $12,200 0.79 $12,200 P-704 1 1 Ammonia Pump 101 367,437 367,437 1.00 $5,000 1997 $10,000 0.79 $10,000 P-705 1 Hydrazine Pump 101 367,437 367,437 1.00 $5,500 1997 $5,500 0.79 $5,500 S-701 1 1 Instrument Air Dryer 101 367,437 367,437 1.00 $8,349 2002 $16,698 0.6 $16,698 T-701 1 Plant Air Receiver 101 367,437 367,437 1.00 $7,003 2002 $7,003 0.72 $7,003 T-702 1 Firewater Storage Tank 101 367,437 367,437 1.00 $166,100 1997 $166,100 0.51 $166,100 

T-703 1 Diesel Storage Tank 101 367,437 367,437 1.00 $14,400 1997 $14,400 0.51 $14,400 T-704 1 Ammonia Storage Tank 101 367,437 367,437 1.00 $287,300 1997 $287,300 0.72 $287,300 

T-705 1 Olivine Lock Hopper 101 367,437 367,437 1.00 Included in gasification & clean up cost (R-201) T-706 1 MgO Lock Hopper 101 367,437 367,437 1.00 Included in gasification & clean up cost (R-201) T-707 1 Hydrazine Storage Tank 101 367,437 367,437 1.00 $12,400 1997 $12,400 0.93 $12,400 A700 Subtotal $1,381,507 $1,381,942 
Appendix I: Goal Design Summary of Individual Equipment Costs 
Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name Scaling Stream 

Scaling Stream 

Flow (lb/hr or 

btu/hr) 

New Stream 

Flow Size Ratio 

Original Equip 

Cost (per unit) Base Year 

Total Original Equip 

Cost (Req'd & Spare) 

in Base Year 

Scaling 

Exponent 

Scaled Cost in 

Base Year 

C-101 4 Hopper Feeder 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-102 2 Screener Feeder Conveyor 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-103 2 Radial Stacker Conveyor 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-104 2 Dryer Feed Screw Conveyor 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) C-105 2 Gasifier Feed Screw Conveyor 104 208,771 208,771 1.00 Included in feed handling & drying cost (M-104) H-101 1 Flue Gas Cooler / Steam Generator #3 PINCH 1,369,986 1,369,094 1.00 $26,143 2002 $26,143 0.6 $26,133 K-101 2 Flue Gas Blower 112 639,530 639,526 1.00 Included in feed handling & drying cost (M-104) M-101 4 Hydraulic Truck Dump with Scale 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) M-102 2 Hammermill 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) M-103 3 Front End Loaders 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) M-104 2 Rotary Biomass Dryer 101 367,437 367,437 1.00 $3,813,728 2002 $7,627,455 0.75 $7,627,450 

S-101 2 Magnetic Head Pulley 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) S-102 2 Screener 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) S-103 2 Dryer Air Cyclone 111 639,530 639,526 1.00 Included in feed handling & drying cost (M-104) S-104 2 Dryer Air Baghouse Filter 103 208,771 208,771 1.00 Included in feed handling & drying cost (M-104) T-101 4 Dump Hopper 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) T-102 1 Hammermill Surge Bin 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) T-103 2 Dryer Feed Bin 101 367,437 367,437 1.00 Included in feed handling & drying cost (M-104) T-104 2 Dried Biomass Hopper 104 208,771 208,771 1.00 Included in feed handling & drying cost (M-104) A100 Subtotal $7,653,598 $7,653,583 C-201 1 Sand/ash Conditioner/Conveyor 219 7,380 7,380 1.00 Included in gasification & clean up cost (R-201) H-201 1 Post-tar Reformer Cooler / Steam Generator #1 PINCH 47,912,711 116,732,109 1.00 $69,089 2002 $69,089 0.65 $69,060 

H-202A 1 Post-tar reformer cooler/Deaerator water preheater #1 PINCH 8,807,704 8,807,704 1.00 $21,589 2002 $21,589 0.6 $21,589 H-202B 1 Post-tar cracker cooler/BFW preheater #2 PINCH 48,632,640 48,632,640 1.00 $429,889 2002 $429,889 0.6 $429,889 K-201 2 Combustion Air Blower 208 442,163 442,157 1.00 Included in gasification & clean up cost (R-201) K-202 1 Regenerator Combustion Air Blower 430 304,578 302,225 0.99 $35,020 2002 $35,020 0.59 $34,860 

M-201 2 Sand/ash Cooler 217 6,642 6,642 1.00 Included in gasification & clean up cost (R-201) R-201 2 Indirectly-heated Biomass Gasifier 201 5,228,880 5,228,878 1.00 $3,318,302 2002 $6,636,603 0.65 $6,636,601 R-202 2 Char Combustor 210 5,434,490 5,434,489 1.00 Included in gasification & clean up cost (R-201) R-203 1 Tar Reformer 225 241,995 241,993 1.00 Included in gasification & clean up cost (R-201) 

R-204 1 Tar Reformer Catalyst Regenerator 428 234,433 234,433 1.00 $2,429,379 2002 $2,429,379 0.65 $2,429,380 

S-201 2 Primary Gasifier Cyclone 202 5,228,880 5,228,878 1.00 Included in gasification & clean up cost (R-201) S-202 2 Secondary Gasifier Cyclone 222 246,484 246,481 1.00 Included in gasification & clean up cost (R-201) S-203 2 Primary Combustor Cyclone 210 5,434,490 5,434,489 1.00 Included in gasification & clean up cost (R-201) S-204 2 Secondary Combustor Cyclone 212 487,506 487,502 1.00 Included in gasification & clean up cost (R-201) S-205 2 Electrostatic Precipitator 213 480,870 480,866 1.00 Included in gasification & clean up cost (R-201) S-206 1 Tar Reformer Cyclone 225 241,995 241,993 1.00 Included in tar reformer catalyst regenerator cost S-207 1 Catalyst Regenerator Cyclone 428 234,433 234,433 1.00 Included in tar reformer catalyst regenerator cost T-201 1 Sand/ash Bin 217 6,642 6,642 1.00 Included in gasification & clean up cost (R-201) A200 Subtotal $9,621,569 $9,621,380 
Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name Scaling Stream 

Scaling Stream 

Flow (lb/hr or 

btu/hr) 

New Stream 

Flow Size Ratio 

Original Equip 

Cost (per unit) Base Year 

Total Original Equip 

Cost (Req'd & Spare) 

in Base Year 

Scaling 

Exponent 

Scaled Cost in 

Base Year 

H-301 1 Quench Water Recirculation Cooler 301 241,995 241,993 1.00 Included in gasification & clean up cost (R-201) H-302 5 Syngas Compressor Intercoolers 301 241,995 241,993 1.00 Included in the syngas compressor cost (K-301) H-303 1 Water-cooled Aftercooler QCH303CT 2,938,799 3,388,287 1.15 $20,889 2002 $20,889 0.44 $22,239 H-304 1 LO-CAT Preheater PINCH 770,434 858,449 1.11 $4,743 2002 $4,743 0.6 $5,061 

H-305 1 LO-CAT Absorbent Solution Cooler 320 179,394 184,842 1.03 Included in LO-CAT oxidizer vessel cost (R-301) H-306 1 ZnO Bed Preheater PINCH 47,209,942 51,594,124 1.09 $71,389 2002 $71,389 0.44 $74,234 

K-301 1 Syngas Compressor 315 220,009 233,488 1.06 $4,817,834 2002 $4,817,834 0.8 $5,052,554 

K-302 1 LO-CAT Feed Air Blower 322 359 358 1.00 Included in LO-CAT oxidizer vessel cost (R-301) K-303 1 Regenerator Flue Gas Blower 434 534,677 536,658 1.00 $54,250 2002 $54,250 0.59 $54,368 

M-301 1 Syngas Quench Chamber 301 241,496 241,993 1.00 Included in gasification & clean up cost (R-201) M-302 1 Syngas Venturi Scrubber 301 241,496 241,993 1.00 Included in gasification & clean up cost (R-201) M-303 1 LO-CAT Venturi Precontactor 323 517 515 1.00 Included in LO-CAT oxidizer vessel cost (R-301) M-304 1 LO-CAT Liquid-filled Absorber 320 179,394 184,842 1.03 Included in LO-CAT oxidizer vessel cost (R-301) P-301 1 1 Sludge Pump 336 997 997 1.00 $3,911 2002 $7,822 0.33 $7,823 P-302 1 1 Quench Water Recirculation Pump 307 1,272,120 316,851 0.25 Included in gasification & clean up cost (R-201) P-303 1 1 LO-CAT Absorbent Solution Circulating Pump 301 241,496 241,993 1.00 Included in LO-CAT oxidizer vessel cost (R-301) R-301 1 LO-CAT Oxidizer Vessel 323 517 515 1.00 $1,000,000 2002 $1,000,000 0.65 $997,471 R-302 2 ZnO Sulfur Removal Beds 327 179,237 184,685 1.03 $37,003 2002 $74,006 0.56 $75,257 

S-301 1 Pre-compressor Knock-out 315 220,009 233,488 1.06 $157,277 2002 $157,277 0.6 $162,989 

S-302 4 Syngas Compressor Interstage Knock-outs 315 220,009 233,488 1.06 Included in the syngas compressor cost (K-301) S-303 1 Post-compressor Knock-out 319 179,394 184,842 1.03 $40,244 2002 $40,244 0.6 $40,973 

T-301 1 Sludge Settling Tank 302 21,718 8,171 0.38 $11,677 2002 $11,677 0.6 $6,495 

T-302 1 Quench Water Recirculation Tank 301 241,496 241,993 1.00 Included in gasification & clean up cost (R-201) A300 Subtotal $6,260,131 $6,499,465 H-404A 1 Tar reformer flue gas cooler/steam generator #2 PINCH 86,510,197 86,510,197 1.00 $144,489 2002 $144,489 0.6 $144,489 

H-404B 1 Tar reformer flue gas cooler/ steam superheater #1 PINCH 108,355,680 108,355,680 1.00 $90,889 2002 $90,889 0.6 $90,889 

H-405 1 LT Shift Precooler/BFW Preheater #1 PINCH 54,476,359 85,423,190 1.57 $56,089 2002 $56,089 0.6 $73,468 

H-407 1 PSA Precooler / Deaerator Water Preheater #2 PINCH 21,034,730 31,414,870 1.49 $21,089 2002 $21,089 0.6 $26,827 

H-408 1 PSA Air-cooled Precooler QAH408 149,281,592 106,741,857 0.72 $388,064 2002 $388,064 0.6 $317,322 

H-409 1 PSA Water-cooled Precooler QCH409CT 8,414,338 7,346,116 0.87 $35,689 2002 $35,689 0.44 $33,619 R-402 1 High Temperature Shift Reactor 404 354,424 322,868 0.91 $465,907 2002 $465,907 0.56 $442,202 R-403 1 Low Temperature Shift Reactor 407 354,424 322,870 0.91 $323,464 2002 $323,464 0.56 $307,007 

S-401 1 Pre-PSA Knock-out #1 413 354,424 322,870 0.91 $129,979 2002 $129,979 0.6 $122,907 

S-402 1 Pre-PSA Knock-out #2 419 242,691 246,017 1.01 $55,291 2002 $55,291 0.6 $55,744 

S-403 1 Pressure Swing Adsorption Unit 424 14,260 15,322 1.07 $4,855,471 2002 $4,855,471 0.6 $5,069,390 

A400 Subtotal $6,566,421 $6,683,865 H-501A 1 Hydrogen Compressor Intercooler QAK501A 4,042,813 4,356,835 1.08 $53,601 2002 $53,601 0.6 $56,062 

H-501B 1 Hydrogen Compressor Air-cooled Aftercooler QAK501B 5,984,714 6,430,563 1.07 $56,901 2002 $56,901 0.6 $59,408 

H-502 1 Hydrogen Compressor Water-cooler Aftercooler QCH502CT 1,465,277 1,574,438 1.07 $18,909 2002 $18,909 0.44 $19,516 K-501 1 Hydrogen Compressor 501 14,260 15,322 1.07 $914,235 2002 $914,235 0.8 $968,331 S-501 1 Pre-hydrogen Compressor Knock-out 501 14,260 15,322 1.07 $13,377 2002 $13,377 0.6 $13,966 

S-502 1 Hydrogen Compressor 1st Interstage Knock-out 501 14,260 15,322 1.07 Included in the hydrogen compressor cost (K-501) S-503 1 Post-hydrogen Compressor Knock-out 505 14,260 15,322 1.07 $13,977 2002 $13,977 0.6 $14,593 

A500 Subtotal $1,071,000 $1,131,876 
Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name Scaling Stream 

Scaling Stream 

Flow (lb/hr or 

btu/hr) 

New Stream 

Flow Size Ratio 

Original Equip 

Cost (per unit) Base Year 

Total Original Equip 

Cost (Req'd & Spare) 

in Base Year 

Scaling 

Exponent 

Scaled Cost in 

Base Year 

H-601 1 Steam Turbine Condenser 614 93,974 131,510 1.40 Included in the extraction steam trubine/generator cost (M-602) H-602 1 Blowdown Cooler / Deaerator Water Preheater PINCH 2,877,029 2,881,506 1.00 $3,043 2002 $3,043 0.6 $3,046 H-603 1 Blowdown Water-cooled Cooler QCH603CT 626,343 627,318 1.00 $16,143 2002 $16,143 0.44 $16,154 

M-601 1 Hot Process Water Softener System 631 349,266 349,809 1.00 $1,031,023 1999 $1,031,023 0.82 $1,032,338 

M-602 1 Extraction Steam Turbine/Generator 607 342,283 342,816 1.00 $4,045,870 2002 $4,045,870 0.71 $4,050,339 

M-603 1 Startup Boiler 200 36,560 36,560 1.00 $198,351 2002 $198,351 0.6 $198,351 

P-601 1 1 Collection Pump 625 255,292 218,299 0.86 $7,015 2002 $14,030 0.33 $13,324 

P-602 1 1 Condensate Pump 616 93,974 131,510 1.40 $5,437 2002 $10,874 0.33 $12,149 

P-603 1 1 Deaerator Feed Pump 628 349,266 349,809 1.00 $8,679 2002 $17,358 0.33 $17,367 

P-604 1 1 Boiler Feed Water Pump 639 349,268 349,812 1.00 $95,660 2002 $191,320 0.33 $191,418 

T-601 1 Condensate Collection Tank 627 349,266 349,809 1.00 $24,493 2002 $24,493 0.6 $24,516 

T-602 1 Condensate Surge Drum 638 349,268 349,812 1.00 $28,572 2002 $28,572 0.6 $28,599 

T-603 1 Deaerator 633 349,266 349,809 1.00 $130,721 2002 $130,721 0.72 $130,867 

T-604 1 Steam Drum 644 349,268 349,812 1.00 $9,200 1997 $9,200 0.72 $9,210 

S-601 1 Blowdown Flash Drum 604 6,985 6,996 1.00 $14,977 2002 $14,977 0.6 $14,991 

A600 Subtotal $5,735,975 $5,742,670 K-701 2 1 Plant Air Compressor 101 367,437 367,437 1.00 $32,376 2002 $97,129 0.34 $97,129 

M-701 1 Cooling Tower System QCTOTAL 139,850,763 145,159,707 1.04 $267,316 2002 $267,316 0.78 $275,199 

M-702 1 Hydraulic Truck Dump with Scale 101 367,437 367,437 1.00 $80,000 1998 $80,000 0.6 $80,000 

M-703 1 Flue Gas Stack 112 1,174,206 639,526 1.00 $51,581 2002 $51,581 1 $51,668 

434 536,658 The stack flow is the sum of two flow streams. 

P-701 1 1 Cooling Water Pump 715 6,088,320 6,319,444 1.04 $158,540 2002 $317,080 0.33 $321,003 

P-702 1 1 Firewater Pump 101 367,437 367,437 1.00 $18,400 1997 $36,800 0.79 $36,800 

P-703 1 1 Diesel Pump 101 367,437 367,437 1.00 $6,100 1997 $12,200 0.79 $12,200 

P-704 1 1 Ammonia Pump 101 367,437 367,437 1.00 $5,000 1997 $10,000 0.79 $10,000 

P-705 1 Hydrazine Pump 101 367,437 367,437 1.00 $5,500 1997 $5,500 0.79 $5,500 

S-701 1 1 Instrument Air Dryer 101 367,437 367,437 1.00 $8,349 2002 $16,698 0.6 $16,698 

T-701 1 Plant Air Receiver 101 367,437 367,437 1.00 $7,003 2002 $7,003 0.72 $7,003 

T-702 1 Firewater Storage Tank 101 367,437 367,437 1.00 $166,100 1997 $166,100 0.51 $166,100 

T-703 1 Diesel Storage Tank 101 367,437 367,437 1.00 $14,400 1997 $14,400 0.51 $14,400 

T-704 1 Ammonia Storage Tank 101 367,437 367,437 1.00 $287,300 1997 $287,300 0.72 $287,300 

T-705 1 Olivine Lock Hopper 101 367,437 367,437 1.00 Included in gasification & clean up cost (R-201) T-706 1 MgO Lock Hopper 101 367,437 367,437 1.00 Included in gasification & clean up cost (R-201) T-707 1 Hydrazine Storage Tank 101 367,437 367,437 1.00 $12,400 1997 $12,400 0.93 $12,400 

A700 Subtotal $1,381,507 $1,393,399 
REPORT DOCUMENTATION PAGE Form Approved 

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PLEASE DO NOT RETURN YOUR FORM TO THE ABOVE ORGANIZATION. 

1. REPORT DATE (DD-MM-YYYY) 

May 2005 

2. REPORT TYPE 

Technical Report 

3. DATES COVERED (From - To) 

4. TITLE AND SUBTITLE 

Biomass to Hydrogen Production Detailed Design and 

Economics Utilizing the Battelle Columbus Laboratory 

Indirectly-Heated Gasifier 

5a. CONTRACT NUMBER 

DE-AC36-99-GO10337 

5b. GRANT NUMBER 

5c. PROGRAM ELEMENT NUMBER 

6. AUTHOR(S) 

P. Spath, A. Aden, T. Eggeman, M. Ringer, B. Wallace, 

and J. Jechura 

5d. PROJECT NUMBER 

NREL/TP-510-37408 

5e. TASK NUMBER 

BB053710 

5f. WORK UNIT NUMBER 

7. PERFORMING ORGANIZATION NAME(S) AND ADDRESS(ES) 

National Renewable Energy Laboratory 

1617 Cole Blvd. 

Golden, CO 80401-3393 

8. PERFORMING ORGANIZATION 

REPORT NUMBER 

NREL/TP-510-37408 

9. SPONSORING/MONITORING AGENCY NAME(S) AND ADDRESS(ES) 10. SPONSOR/MONITOR'S ACRONYM(S) 

NREL 

11. SPONSORING/MONITORING 

AGENCY REPORT NUMBER 

12. DISTRIBUTION AVAILABILITY STATEMENT 

National Technical Information Service 

U.S. Department of Commerce 

5285 Port Royal Road 

Springfield, VA 22161 

13. SUPPLEMENTARY NOTES 

14. ABSTRACT (Maximum 200 Words) 

This analysis developed detailed process flow diagrams and an Aspen Plus® model, evaluated energy flows 

including a pinch analysis, obtained process equipment and operating costs, and performed an economic evaluation 

of two process designs based on the syngas clean up and conditioning work being performed at NREL. One design, 

the current design, attempts to define today’s state of the technology. The other design, the goal design, is a target 

design that attempts to show the effect of meeting specific research goals. 

15. SUBJECT TERMS 

Battelle; Columbus; gasifier; energy flow; economic evaluation 

16. SECURITY CLASSIFICATION OF: 17. LIMITATION 

OF ABSTRACT 

UL 

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