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

Efficiency & Renewable Energy, operated by the Alliance for Sustainable Energy, LLC.

Contract No. DE-AC36-08GO28308

Process Design and Economics 

for Biochemical Conversion of 

Lignocellulosic Biomass to 

Ethanol 

Dilute-Acid Pretreatment and 

Enzymatic Hydrolysis of Corn Stover

D. Humbird, R. Davis, L. Tao, C. Kinchin, 

D. Hsu, and A. Aden

National Renewable Energy Laboratory

Golden, Colorado

P. Schoen, J. Lukas, B. Olthof, M. Worley, 

D. Sexton, and D. Dudgeon

Harris Group Inc.

Seattle, Washington and Atlanta, Georgia

Technical Report

NREL/TP-5100-47764

May 2011
NREL is a national laboratory of the U.S. Department of Energy, Office of Energy 

Efficiency & Renewable Energy, operated by the Alliance for Sustainable Energy, LLC.

National Renewable Energy Laboratory

1617 Cole Boulevard

Golden, Colorado 80401

303-275-3000 • www.nrel.gov

Contract No. DE-AC36-08GO28308

Process Design and Economics 

for Biochemical Conversion of 

Lignocellulosic Biomass to 

Ethanol 

Dilute-Acid Pretreatment and 

Enzymatic Hydrolysis of Corn Stover

D. Humbird, R. Davis, L. Tao, C. Kinchin, 

D. Hsu, and A. Aden

National Renewable Energy Laboratory

Golden, Colorado

P. Schoen, J. Lukas, B. Olthof, M. Worley, 

D. Sexton, and D. Dudgeon

Harris Group Inc.

Seattle, Washington and Atlanta, Georgia

Prepared under Task No. BB07.2410

Technical Report

NREL/TP-5100-47764

May 2011
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iii

Executive Summary

The U.S. Department of Energy (DOE) promotes the production of ethanol and other liquid fuels 

from lignocellulosic biomass feedstocks by funding fundamental and applied research that 

advances the state of technology in biomass collection, conversion, and sustainability. As part of 

its involvement in the program, the National Renewable Energy Laboratory (NREL) investigates 

the production economics of these fuels. 

This report describes in detail one potential biochemical ethanol conversion process, 

conceptually based upon core conversion and process integration research at NREL. The 

overarching process design converts corn stover to ethanol by dilute-acid pretreatment, 

enzymatic saccharification, and co-fermentation. Ancillary areas—feed handling, product 

recovery, wastewater treatment, lignin combustion, and utilities—are also included in the design. 

Detailed material and energy balances and capital and operating costs were developed for the 

entire process, and they are documented in this report and accompanying process simulation 

files, which are available to the public.

As a benchmark case study, this so-called technoeconomic model provides an absolute 

production cost for ethanol that can be used to assess its competitiveness and market potential. It 

can also be used to quantify the economic impact of individual conversion performance targets 

and prioritize these in terms of their potential to reduce cost. Furthermore, by using the 

benchmark as a comparison, DOE can make more informed decisions about research proposals

claiming to lower ethanol production costs. 

Building on design reports published in 2002 and 1999, NREL, together with the subcontractor 

Harris Group Inc., performed a complete review of the process design and economic model for 

the biomass-to-ethanol process. This update reflects NREL’s current vision of the biochemical 

ethanol process and incorporates recent progress in the conversion areas (pretreatment, 

conditioning, saccharification, and fermentation), optimizations in product recovery, and an 

improved understanding of the ethanol plant’s back end (wastewater and utilities). The major 

process updates in this design report are the following:

• Feedstock composition is updated to a carbohydrate profile closer to the expected mean. 

• Pretreatment reactor configuration is revised with significant new detail. 

• Whole-slurry pH adjustment of the pretreated biomass with ammonia replaced the 

previous conditioning practice of overliming, eliminating a solid-liquid separation step. 

• Enzymatic hydrolysis and fermentation are modeled as a batch process with a continuous 

high-solids hydrolysis reactor upstream of the batch reactors. 

• On-site enzyme production is included to increase transparency on the cost of enzymes. 

• Wastewater treatment section is redesigned to handle inorganics in the ethanol stillage. 

The conceptual design presented here reports ethanol production economics as determined by 

2012 conversion targets and “n

th-plant” project costs and financing. For the biorefinery described 

here, processing 2,205 dry ton/day at 76% theoretical ethanol yield (79 gal/dry ton), the ethanol 

selling price is $2.15/gal in 2007$. 
iv

Figure ES-1. Economic summary for ethanol production

Minimum Ethanol Selling Price (MESP): $2.15 /gal

Gasoline-Equivalent MESP: $3.27 /gal gasoline equivalent

Contributions: Feedstock $0.74 /gal

Enzymes $0.34 /gal

Non-Enzyme Conversion $1.08 /gal

Ethanol Production 61.0 MMgal/yr (Ethanol at 68 °F)

Ethanol Yield 79.0 gal / dry U.S. ton feedstock

Feedstock + Handling Cost $58.50 /dry U.S. ton

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 40%

Capital Costs Manufacturing Costs (cents/gal ethanol)

Pretreatment $29,900,000 Feedstock + Handling 74.1

Neutralization/Conditioning $3,000,000 Sulfuric Acid 2.4

Saccharification & Fermentation $31,200,000 Ammonia 6.5

On-site Enzyme Production $18,300,000 Glucose (enzyme production) 19.3

Distillation and Solids Recovery $22,300,000 Other Raw Materials 12.9

Wastewater Treatment $49,400,000 Waste Disposal 2.5

Storage $5,000,000 Net Electricity -10.8

Boiler/Turbogenerator $66,000,000 Fixed Costs 17.5

Utilities $6,900,000 Capital Depreciation 22.0

Total Installed Equipment Cost $232,000,000 Average Income Tax 12.3

Average Return on Investment 56.6

Added Direct + Indirect Costs $190,500,000

 (% of TCI) 45% Manufacturing Costs ($/yr)

Feedstock + Handling $45,200,000

Total Capital Investment (TCI) $422,500,000 Sulfuric Acid $1,500,000

Ammonia $4,000,000

Installed Equipment Cost/Annual Gallon $3.80 Glucose (enzyme production) $11,800,000

Total Capital Investment/Annual Gallon $6.92 Other Raw Materials $7,900,000

Waste Disposal $1,500,000

Loan Rate 8.0% Net Electricity -$6,600,000

Term (years) 10 Fixed Costs $10,700,000

Capital Charge Factor (Computed) 0.131 Capital Depreciation $13,400,000

Average Income Tax $7,500,000

Denatured Fuel Production (MMgal/yr) 61.7 Average Return on Investment $34,600,000

Denatured Fuel Min. Sales Price $2.18

Denaturant Cost ($/gal denaturant) $2.10 Specific Operating Conditions

Enzyme Loading (mg/g cellulose) 20

Maximum Yields (100% of Theoretical) Saccharification Time (days) 3.5

 Ethanol Production (MMgal/yr) 80.3 Fermentation Time (days) 1.5

 Theoretical Yield (gal/U.S. ton) 103.9 Ethanol titer (wt%) 5.4%

Current Yield (Actual/Theoretical) 76.0% Excess Electricity (kWh/gal) 1.8

Plant Electricity Use (kWh/gal) 3.9

Plant Water Usage (gal/gal) 5.4

All Values in 2007$

Dilute Acid Pretreatment with Enzymatic Hydrolysis and Co-Fermentation

Corn Stover Design Report Case: 2012 model DW1102A

Ethanol Production Process Engineering Analysis
v 

Table of Contents

1 Introduction........................................................................................................................................... 1

1.1 Background and Motivation ................................................................................................1

1.2 Process Overview .................................................................................................................2

1.3 Technoeconomic Analysis Approach ..................................................................................5

1.4 About n

th-Plant Assumptions ...............................................................................................6

1.5 Review of Related Technoeconomic Studies ......................................................................7

1.6 About the NREL Aspen Model..........................................................................................10

2 Design Basis and Conventions......................................................................................................... 12

2.1 Plant Size ...........................................................................................................................12

2.2 Feedstock Composition ......................................................................................................12

2.3 Design Report Conventions ...............................................................................................15

2.3.1 Units ..........................................................................................................................15

2.3.2 Total Solids Loading .................................................................................................15

2.3.3 Ethanol Density .........................................................................................................15

2.3.4 Theoretical Yields and Conversions .........................................................................15

3 Process Design and Cost Estimation Details.................................................................................. 16

3.1 Area 100: Feedstock Storage and Handling .......................................................................16

3.1.1 Overview ...................................................................................................................16

3.1.2 Design Basis ..............................................................................................................17

3.1.3 Cost Estimation .........................................................................................................17

3.2 Area 200: Pretreatment and Conditioning .........................................................................19

3.2.1 Overview ...................................................................................................................19

3.2.2 Design Basis ..............................................................................................................20

3.2.3 Cost Estimation .........................................................................................................22

3.2.4 Achieving the Design Case .......................................................................................23

3.3 Area 300: Enzymatic Hydrolysis and Fermentation ..........................................................26

3.3.1 Overview ...................................................................................................................26

3.3.2 Design Basis ..............................................................................................................27

3.3.3 Cost Estimation .........................................................................................................31

3.3.4 Achieving the Design Case .......................................................................................32

3.4 Area 400: Cellulase Enzyme Production ...........................................................................36

3.4.1 Overview ...................................................................................................................36

3.4.2 Design Basis ..............................................................................................................37

3.4.3 Cost Estimation .........................................................................................................40

3.4.4 Enzyme Cost Discussion ...........................................................................................40

3.5 Area 500: Product, Solids, and Water Recovery ...............................................................44

3.5.1 Overview ...................................................................................................................44

3.5.2 Design Basis ..............................................................................................................45

3.5.3 Cost Estimation .........................................................................................................46

3.6 Area 600: Wastewater Treatment (WWT) .........................................................................47

3.6.1 Overview ...................................................................................................................47

3.6.2 Design Basis ..............................................................................................................48

3.6.3 Cost Estimation .........................................................................................................50

3.6.4 Wastewater Challenges .............................................................................................50

3.7 Area 700: Product and Feed Chemical Storage .................................................................51
vi 

3.7.1 Overview ...................................................................................................................51

3.7.2 Design Basis ..............................................................................................................51

3.7.3 Cost Estimation .........................................................................................................51

3.8 Area 800: Combustor, Boiler, and Turbogenerator ...........................................................52

3.8.1 Overview ...................................................................................................................52

3.8.2 Design Basis ..............................................................................................................52

3.8.3 Cost Estimation .........................................................................................................53

3.9 Area 900: Utilities ..............................................................................................................54

3.9.1 Overview ...................................................................................................................54

3.9.2 Design Basis ..............................................................................................................54

3.9.3 Cost Estimation .........................................................................................................56

4 Process Economics ........................................................................................................................... 57

4.1 About Cost-Year Indices ....................................................................................................57

4.2 Total Capital Investment (TCI) ..........................................................................................58

4.3 Variable Operating Costs ...................................................................................................62

4.4 Fixed Operating Costs........................................................................................................64

4.5 Discounted Cash Flow Analysis and the Minimum Selling Price of Ethanol ...................65

5 Analysis and Discussion ................................................................................................................... 70

5.1 Carbon and Energy Balance...............................................................................................70

5.2 Water Balance ....................................................................................................................74

5.3 The Cost of Sugar ..............................................................................................................76

5.4 Cost Sensitivity Analysis ...................................................................................................79

5.5 State of Technology Back-Casting ....................................................................................81

6 Concluding Remarks.......................................................................................................................... 83

6.1 Summary ............................................................................................................................83

6.2 Variance from the 2002 Design .........................................................................................84

6.3 Future Work .......................................................................................................................85

References................................................................................................................................................. 87

Appendix A. Individual Equipment Costs Summary ............................................................................. 96

Appendix B. Discounted Cash Flow Rate of Return Worksheet........................................................ 103

Appendix C. Process Parameters/Operating Summary...................................................................... 106

Appendix D. Aspen Properties .............................................................................................................. 107

Appendix E. Process Flow Diagrams.................................................................................................... 114
vii 

List of Figures

Figure 1. Simplified flow diagram of the overall process, PFD-P120-A000 ................................. 4

Figure 2. NREL’s approach to process design and economic analysis .......................................... 5

Figure 3. Adjusting the Kazi et al. base case ($3.40/gal) to NREL 2008 SOT assumptions ........ 10

Figure 4. Structural carbohydrate distribution of the feedstock variability study [25] ................. 13

Figure 5. Simplified flow diagram of the feedstock receiving system ......................................... 16

Figure 6. Simplified flow diagram of the pretreatment and conditioning process ....................... 19

Figure 7. Horizontal pretreatment reactor design scheme (Andritz, Inc.) .................................... 23

Figure 8. Simplified flow diagram of the enzymatic hydrolysis and fermentation process ......... 27

Figure 9. Simplified flow diagram of the enzyme production process ......................................... 36

Figure 10. Enzyme production cost breakdown ........................................................................... 41

Figure 11. Enzyme contribution to MESP as a function of loading ............................................. 42

Figure 12. Simplified flow diagram of the separation process ..................................................... 44

Figure 13. Simplified flow diagram of the wastewater treatment process ................................... 47

Figure 14. Cooling duty distribution between major users ........................................................... 55

Figure 15. Distribution of plant electricity utilization by process area ......................................... 56

Figure 16. Sensitivity of MESP to IRR and % equity (8% interest on a 10-year loan) ................ 66

Figure 17. Cost contribution details from each process area (per gallon EtOH) .......................... 69

Figure 18. Mass, carbon, and energy balance diagram for the overall process ............................ 73

Figure 19. Economic summary for dilute sugar production ......................................................... 78

Figure 20. Single-point sensitivity tornado charts for (a) MESP and (b) ethanol yield ............... 80

Figure 21. Sugar-based pathways for conventional and advanced biofuels ................................. 86
viii

List of Tables

Table 1. Comparison of Technoeconomic Analyses and Market Studies for Existing Biofuels .... 7

Table 2. Survey of Recent Technoeconomic Studies on Biochemical Cellulosic Ethanol ............ 9

Table 3. Summary of Whole Stover Composition Data ............................................................... 13

Table 4. Corn Stover Composition from the 2002 Design [2] and the Present Design ................ 14

Table 5. Pretreatment Reactor Conditions .................................................................................... 21

Table 6. Pretreatment Hydrolysis Reactions and Assumed Conversions ..................................... 22

Table 7. Research Status and 2012 Targets in the Pretreatment Area .......................................... 24

Table 8. Enzymatic Hydrolysis Conditions .................................................................................. 28

Table 9. Enzymatic Hydrolysis Reactions and Assumed Conversions ........................................ 28

Table 10. Seed Train Specifications ............................................................................................. 29

Table 11. Seed Train Reactions and Assumed Conversions ......................................................... 29

Table 12. Co-Fermentation Contamination Loss Reactions ......................................................... 30

Table 13. Co-Fermentation Conditions......................................................................................... 31

Table 14. Co-Fermentation Reactions and Assumed Conversions ............................................... 31

Table 15. Research Status and 2012 Targets in Enzymatic Hydrolysis and Fermentation .......... 32

Table 16. Summary of Integrated Whole-Slurry Enzymatic Hydrolysis Performance Using an 

Advanced Enzyme Preparation ..................................................................................................... 33

Table 17. Ethanol Yields from Xylose ......................................................................................... 34

Table 18. Area 400 Guiding Design Basis Assumptions .............................................................. 37

Table 19. Specifications of the Enzyme Production Bioreactors ................................................. 39

Table 20. Cellulase Production Nutrient Requirements ............................................................... 39

Table 21. Cellulase Production Electricity Requirements ............................................................ 40

Table 22. Wastewater Treatment System Design Basis ............................................................... 48

Table 23. Storage Requirements ................................................................................................... 51

Table 24. Cooling Water Users ..................................................................................................... 54

Table 25. Installation Factors ........................................................................................................ 59

Table 26. Scaling Exponents ......................................................................................................... 60

Table 27. Additional Costs for Determining Total Capital Investment (TCI) .............................. 61

Table 28. Project Cost Worksheet Including Total Direct Costs and Total Capital Investment .. 62

Table 29. Chemical Costs and Sources ......................................................................................... 63

Table 30. Variable Operating Costs .............................................................................................. 63

Table 31. Fixed Operating Costs .................................................................................................. 64

Table 32. Construction Activities and Cash Flow ........................................................................ 67

Table 33. Discounted Cash Flow Analysis Parameters ................................................................ 68

Table 34. Summary of Yields, Rates, and Conversion Costs ....................................................... 69

Table 35. Ethanol Plant Overall Carbon Balance ......................................................................... 70

Table 36. Ethanol Plant Overall Energy Balance ......................................................................... 71

Table 37. Ethanol Plant Water Balance ........................................................................................ 75

Table 38. Individual Contributors to Cooling Water Evaporation ............................................... 75

Table 39. Mixed Sugar Stream Composition ................................................................................ 77

Table 40. Assumptions Varied in the Sensitivity Analysis ........................................................... 79

Table 41. Back-Casted State of Technology Cases and Future Targets ....................................... 82

Table 42. Comparison of the Previous Target Case and the Present Design (2007$/gal) ............ 84

Table 43. Individual Contributors to the Change in MESP .......................................................... 85
ix

Acronyms

AFEX ammonia fiber explosion 

BFW boiler feed water

BLS Bureau of Labor Statistics

C5/C6 refers to mixtures of xylose (a C5 sugar) and glucose (a C6 sugar)

CBP consolidated bioprocessing 

CIP clean-in-place 

COD chemical oxygen demand

CSL corn steep liquor 

DAP diammonium phosphate 

DCFROR discounted cash flow rate of return

DOE U.S. Department of Energy 

FCI fixed capital investment 

FGD flue gas desulfurization

HHV higher heating value 

HMF 5-hydroxymethyl furfural 

INL Idaho National Laboratory

IRR internal rate of return

ISBL inside battery limits (of the plant) 

LHV lower heating value 

MESP minimum ethanol selling price

MM million (e.g., MMBtu or $__MM)

MSSP minimum sugar selling price 

MYPP OBP’s Multi-Year Program Plan

NPV net present value 

NREL National Renewable Energy Laboratory 

OBP Office of the Biomass Program

OTR oxygen transfer rate

OUR oxygen uptake rate

PCS pretreated corn stover 

PFD process flow diagram

SCFM standard cubic feet per minute 

SHF separate (or sequential) hydrolysis and fermentation

SOT annual State of Technology case 

SSCF simultaneous saccharification and co-fermentation

TCI total capital investment 

TDC total direct cost 

TE technoeconomic 

VOC volatile organic compound 

VVM volume (of gas) per volume (of liquid) per minute 

WWT wastewater treatment

 
1 

1 Introduction

1.1 Background and Motivation

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

production of ethanol and other liquid fuels from lignocellulosic feedstocks by sponsoring 

programs in fundamental and applied research that aim to advance the state of biomass 

conversion technology. These programs include laboratory campaigns to develop better cellulose 

hydrolysis enzymes and fermenting microorganisms, detailed engineering studies of potential 

processes, and construction of pilot-scale demonstration and production facilities. This research 

is conducted by national laboratories, universities, and private industry in conjunction with 

engineering and construction companies. 

As part of its involvement in the program, the National Renewable Energy Laboratory (NREL) 

investigates the complete process design and economics of cellulosic ethanol manufacturing in 

order to develop an absolute plant-gate price for ethanol based on process and plant design 

assumptions consistent with applicable best practices in engineering, construction, and operation. 

This plant-gate price is referred to as the minimum ethanol selling price or MESP. The MESP 

can be used by policymakers and DOE to assess the cost-competitiveness and market penetration 

potential of cellulosic ethanol in comparison with petroleum-derived fuels and starch- or sugarbased ethanol.

The technoeconomic analysis effort at NREL also helps to direct our biomass conversion 

research by examining the sensitivity of the MESP to process alternatives and research advances. 

Proposed research and its anticipated results can be translated into a new MESP that can be 

compared to the benchmark case documented in this report. Such comparison helps to quantify 

the economic impact of core research targets at NREL and elsewhere and to track progress 

toward meeting competitive cost targets. It also allows DOE to make more informed decisions 

about research proposals that claim to reduce MESP.

This report builds upon previous issues from 1999 [1] and 2002 [2] written by NREL engineers 

with Delta-T, Merrick Engineering, Reaction Engineering, Inc., and Harris Group. For the 

present report, NREL again contracted Harris Group to provide engineering support for 

estimating and reviewing the equipment and raw material costs used in the process design. This 

update reflects NREL’s latest envisioned biochemical ethanol process and includes recent 

research progress in the conversion areas (pretreatment, conditioning, enzymatic hydrolysis, and 

fermentation), optimizations in product recovery, and our latest understanding of the ethanol 

plant’s back end (separation, wastewater, and utilities). NREL worked with Harris Group to 

identify realistic configurations and costs for critical equipment, the pretreatment reactor system

in particular. An on-site cellulase enzyme section was included in this update to permit better 

transparency of enzyme economics than the fixed cost contribution assumed in the last design 

report did. 

The biomass conversion efficiencies used in the design (e.g., cellulose to glucose or xylose to 

ethanol) are based on a slate of research targets that NREL and DOE have committed to 

demonstrate by the end of 2012 in a campaign of integrated pilot-scale runs. These 2012 

performance targets are discussed in detail in this report. The economics of this conceptual 

process use the best available equipment and raw material costs and an “n

th-plant” project cost 
2 

structure and financing. The projected 2012 n

th-plant MESP computed in this report is $2.15/gal 

in 2007$. 

Modifications to the conceptual process design presented here will be reflected annually through 

NREL’s State of Technology (SOT) reports. These ensure that the process design and its cost 

benchmark incorporate the most current data from NREL and other DOE-funded research and 

that equipment costs stay up-to-date. 

We stress that this design report serves to describe a single, feasible cellulosic ethanol 

conversion process and to transparently document the assumptions and details that went into its 

design. This report is not meant to provide an exhaustive survey of process alternatives or costsensitivity analyses. These will be investigated in separate papers that extend and reference the 

present report. Furthermore, the process models and economic tools developed for this report are 

available to the public, and the authors and members of NREL’s Biochemical Platform Analysis 

task will provide support to researchers who wish to use them for their own studies. 

1.2 Process Overview

The process described here uses co-current dilute-acid pretreatment of lignocellulosic biomass 

(corn stover), followed by enzymatic hydrolysis (saccharification) of the remaining cellulose, 

followed by fermentation of the resulting glucose and xylose to ethanol. The process design also 

includes feedstock handling and storage, product purification, wastewater treatment, lignin 

combustion, product storage, and required utilities. The process is divided into nine areas (see

Figure 1).

• Area 100: Feed handling. The feedstock, in this case milled corn stover, is delivered to 

the feed handling area from a uniform-format feedstock supply system. Only minimum 

storage and feed handling are required. From there, the biomass is conveyed to the 

pretreatment reactor (Area 200).

• Area 200: Pretreatment and conditioning. In this area, the biomass is treated with dilute 

sulfuric acid catalyst at a high temperature for a short time to liberate the hemicellulose 

sugars and break down the biomass for enzymatic hydrolysis. Ammonia is then added to 

the whole pretreated slurry to raise its pH from ~1 to ~5 for enzymatic hydrolysis. 

• Area 300: Enzymatic hydrolysis and fermentation. Enzymatic hydrolysis is initiated in a 

high-solids continuous reactor using a cellulase enzyme prepared on-site. The partially 

hydrolyzed slurry is next batched to one of several parallel bioreactors. Hydrolysis is 

completed in the batch reactor, and then the slurry is cooled and inoculated with the cofermenting microorganism Zymomonas mobilis. After a total of five days of sequential 

enzymatic hydrolysis and fermentation, most of the cellulose and xylose have been 

converted to ethanol. The resulting beer is sent to the product recovery train (Area 500). 

• Area 400: Cellulase enzyme production. An on-site enzyme production section was 

included in this design. Purchased glucose (corn syrup) is the primary carbon source for 

enzyme production. Media preparation involves a step in which a portion of the glucose 

is converted to sophorose to induce cellulase production. The enzyme-producing fungus 

(modeled after Trichoderma reesei) is grown aerobically in fed-batch bioreactors. The 
3 

entire fermentation broth, containing the secreted enzyme, is fed to Area 300 to carry out 

enzymatic hydrolysis. 

• Area 500: Product recovery. The beer is separated into ethanol, water, and residual solids 

by distillation and solid-liquid separation. Ethanol is distilled to a nearly azeotropic 

mixture with water and then purified to 99.5% using vapor-phase molecular sieve 

adsorption. Solids recovered from the distillation bottoms are sent to the combustor (Area 

800) while the liquid is sent to wastewater treatment (Area 600).

• Area 600: Wastewater treatment. Plant wastewater streams are treated by anaerobic and 

aerobic digestion. The methane-rich biogas from anaerobic digestion is sent to the 

combustor (Area 800), where sludge from the digesters is also burned. The treated water 

is suitable for recycling and is returned to the process. 

• Area 700: Storage. This area provides bulk storage for chemicals used and produced in 

the process, including corn steep liquor (CSL), ammonia, sulfuric acid, nutrients, water, 

and ethanol. 

• Area 800: Combustor, boiler, and turbogenerator. The solids from distillation and 

wastewater treatment and the biogas from anaerobic digestion are combusted to produce 

high-pressure steam for electricity production and process heat. The majority of the 

process steam demand is in the pretreatment reactor and distillation columns. The boiler 

produces excess steam that is converted to electricity for use in the plant and for sale to 

the grid. 

• Area 900: Utilities. This area includes a cooling water system, chilled water system, 

process water manifold, and power systems.
4 

Figure 1. Simplified flow diagram of the overall process, PFD-P120-A000

A100

FEED HANDLING

A200

PRETREATMENT

& CONDITIONING

A400

ENZYME 

PRODUCTION

EN

HYDR

FER

DIS

DEH

SE

S

A800

BURNER/BOILER 

TURBOGENERATOR

A600

WASTEWATER

TREATMENT

A900

UTILITIES

MA

FEEDSTOCK RAW FEEDSTOCK

RECYCLE 

WATER

NUTRIENTS

AUX. FUEL

ANAEROBIC 

BIOGAS

STEAM

ELECTRICITY

STEAM

CELLULASE

HYDROLYSATE

AMMONIA

STEAM

ACID

GLUCOSE

NUTRIENTS

NUTRIENTS

EVAPORATION FLASH

CONDENSATE

STILLAGE

LIGNIN

VENT
5 

1.3 Technoeconomic Analysis Approach

Figure 2 describes the engineering approach used here for modeling the conversion of biomass to 

ethanol, including process design, process modeling, and economic analysis. 

Figure 2. NREL’s approach to process design and economic analysis

Starting from the general process flow diagram (PFD) shown in Figure 1 and the more detailed 

PFDs contained in Appendix E, a process simulation is developed using Aspen Plus software [3]. 

This process model computes thermodynamically-rigorous material and energy balances for each 

unit operation in this conceptual biorefinery. 

The material and energy balance data from the Aspen simulation are next used to assist in

determining the number and size of capital equipment items. As process conditions and flows 

change, baseline equipment costs are automatically adjusted in an Excel spreadsheet using a 

scaling exponent. These baseline costs come from vendor quotes (favored for larger or nonstandard unit operations and packaged or skid-mounted subsystems) or from Harris Group’s 

Process Flow Diagrams

Rigorous

Material & Energy Balances

(Aspen Plus)

Capital & Project Cost 

Estimation

Discounted Cash Flow 

Economic Model

Minimum Ethanol Selling 

Price

Engineering Companies 

Consulting on Process 

Configuration

Estimates of Other 

Commercial Technology

Cost Estimation 

Software (e.g., ICARUS)

Engineering Company

Cost Estimations

NREL or other

DOE-Sponsored

Research

Vendor Cost Quotations

Outside Engineering

Studies (e.g., feed handling, 

separations)
6 

proprietary cost database (for secondary equipment such as tanks, pumps, and heat exchangers).

Final equipment costs for this report are tabulated in Appendix A. 

Once equipment costs are determined, direct and indirect overhead cost factors (e.g., installation

costs and project contingency) are applied to determine a total capital investment (TCI). The 

TCI, along with the plant operating expenses (also developed using flow rates from the Aspen 

model), is used in a discounted cash flow rate of return (DCFROR) analysis to determine a plantgate price for ethanol for a given discount rate. This plant-gate price is also called the minimum 

ethanol selling price (MESP, $/gallon) required to obtain a net present value (NPV) of zero for a 

10% internal rate of return (IRR) after taxes.

The product of the analysis described above is a technoeconomic model that reasonably 

estimates a product price for a pre-commercial process. The resultant MESP is unique for the set 

of process conditions simulated and it should be emphasized that a certain percentage of 

uncertainty always exists around these chosen conditions, as well as around the assumptions 

made for capital and raw material costs. Without a detailed understanding of the basis behind it, 

the absolute computed MESP has limited relevance. While the MESP can be used to assess the 

marketplace competiveness of a given process, it is best suited for comparing technological 

variations against one another or for performing sensitivity analyses that indicate where 

economic or process performance improvements are needed. 

1.4 About nth-Plant Assumptions

The technoeconomic analysis reported here uses what are known as “n

th-plant” economics. The 

key assumption implied by n

th-plant economics is that our analysis does not describe a pioneer 

plant; instead, several plants using the same technology have already been built and are 

operating. In other words, it reflects a mature future in which a successful industry of n plants 

has been established. Because the technoeconomic model is primarily a tool for studying new 

process technologies or integration schemes in order to comment on their comparative economic 

impact, we feel it is prudent to ignore artificial inflation of project costs associated with risk 

financing, longer start-ups, equipment overdesign, and other costs associated with first-of-a-kind 

or pioneer plants, lest these overshadow the real economic impact of research advances in 

conversion or process integration. At the very least, these n

th-plant economics should help to 

provide justification and support for early technology adopters and pioneer plants. 

In previous design reports that targeted 2012 technology, many of the n

th-plant assumptions also 

applied to key 2012 research targets in conversion, e.g., 90% conversion of xylan to xylose, 

which in 2002 had not yet been experimentally demonstrated. This report will show that through 

research progress, current conversion performance is approaching the targets set in 2002 and that 

the 2012 performance targets should no longer be considered speculative. Assumptions in the 

previous model also extended to enzyme costs and some equipment costs, particularly those 

items that were not commercially available at the time (e.g., the pretreatment reactor). In the 

present design report, enzyme and equipment costs are considerably more well-defined. 

The remaining n

th-plant assumptions in the present model therefore apply primarily to the 

factored cost model used to determine the total capital investment from the purchased equipment 

cost and to the choices made in plant financing. The n

th-plant assumption also applies to some 

operating parameters, such as process uptime of 96%. These assumptions were agreed upon by 
7 

NREL and DOE for this report and reflect our best estimates at the time of publication. It should 

be emphasized, however, that these assumptions carry a large uncertainty and are subject to 

refinement. 

1.5 Review of Related Technoeconomic Studies

Tao and Aden [4] performed a survey of technoeconomic models of existing biofuels (corn 

ethanol, sugarcane ethanol, and conventional soy biodiesel) from the literature. These studies 

were normalized to a consistent year-dollar value and feedstock cost (where applicable) and 

compared to published market studies. This comparison, shown in Table 1, indicated that 

technoeconomic analysis was able to predict the actual cost of production of these biofuels 

within the expected accuracy of such models. 

Table 1. Comparison of Technoeconomic Analyses and Market Studies for Existing Biofuels 

Fuel Market Study TE Model 

Corn ethanol $1.53/gal a $1.54/gal b

Sugarcane ethanol $1.14/gal a $1.29/gal c

Soy biodiesel $2.15/gal d $2.55/gal e

a F.O. Lichts, 2007 [5]. b Kwiatkowski et al., 2006 [6]. c Rodrigues, 2007 [7]; Seabra, 2007 [8]. d Graboski & McCormick, 1998 [9]. e Haas et al., 2006 [10]. 

The quantitative agreement in Table 1 is not altogether surprising, because the developers of 

these models were able to compare their results to real economics and make adjustments as

necessary. (Also, for these particular fuels, feedstock makes up the largest part of the production 

cost, so uncertainty in the remaining non-feedstock conversion costs is not as obvious.) 

Developing a technoeconomic model for a pre-commercial technology, such as the corn-stoverto-ethanol process described in this report, requires a more ab initio approach, rooted in a 

thorough understanding of the state of the technology at the time of the analysis and good 

engineering practice. 

In the years since the last NREL design report in 2002, several newer technoeconomic studies of 

biochemical cellulosic ethanol production have been published. Many of these studies were 

based to varying degrees on NREL’s previous design report, borrowing from its process 

assumptions, cost information, or both. This was in fact the principal goal of the earlier NREL 

design reports: to establish a baseline or “zero-point” technoeconomic benchmark from which 

process alternatives and improvements could be evaluated by others in the public realm. The 

present report now aims to establish a new zero-point and thereby serve a similar purpose for 

analyses going forward. 

A brief survey of MESPs from recent technoeconomic studies of biochemical cellulosic ethanol 

production is presented in Table 2; note that these studies were not normalized to a consistent 

cost-year but were all published between 2008 and 2010, so cost-year differences should be 

minor. Clearly, there is a wide range of published MESP values within this subset of papers. For 

the most part, these are due to differences in feedstock cost, process assumptions, and co-product 

values, all of which vary considerably across the studies. For example, the analysis by Laser et 

al. [11] assumes a fairly low feedstock cost and very high yields (indicating aggressive process 
8 

assumptions) as well as improved economies of scale (if such a high feed rate can be sustained), 

while also receiving positive revenue from higher-value co-products such as protein and 

hydrogen. Furthermore, this study assumes a consolidated bioprocessing (CBP) approach, 

which—although less developed than separate saccharification and co-fermentation—could 

further improve economics by reducing enzyme costs. Conversely, studies on the high end of the 

MESP range such as Kazi et al. [12] and Klein-Marcuschamer et al. [13] assumed higher 

feedstock costs while achieving much lower ethanol yields. Compared to NREL’s 2002 design 

report (upon which many of its cost inputs were based), the Klein-Marcuschamer study’s base 

case assumes much longer batch times for saccharification and fermentation (thus higher 

associated capital costs), higher enzyme costs, and a lower carbohydrate fraction in the feedstock 

(contributing to lower yields) [14].
9 

Table 2. Survey of Recent Technoeconomic Studies on Biochemical Cellulosic Ethanol

Source MESP 

($/gal)

Feedstock Scale

(dry 

ton/day) 

Feed 

Price 

($/dry 

ton)

Ethanol 

Yield 

(gal/dry 

ton)

Notes

NREL 2002 

Design Report 

(2007$) [2]

1.49 Corn stover 2,200 51 90 Dilute-acid 

pretreatment, SSCF 

process, electricity 

co-product

NREL 2011 

Design Report 

(this work)

2.15 Corn stover 2,200 59 79 Dilute-acid 

pretreatment, SHF 

process, electricity 

co-product

Kazi [12] 3.40-

4.44

Corn stover 2,200 75 42-72 Varying pretreatment 

options and 

downstream process 

assumptions

Huang [15] 1.42-

1.87

Aspen, poplar, 

corn stover, 

switchgrass

2,200 58-100 83-111 Dilute-acid 

pretreatment 

Sendich [16] 1.03-

1.41

Corn stover 2,200 40 70 AFEX pretreatment, 

SSCF process, 

varying process 

conditions

Sendich [16] 0.80-

0.95

Corn stover 2,200 40 78 AFEX pretreatment, 

CBP process, varying 

process conditions

Laser [11] 0.63-

0.83

Switchgrass 5,000 44 97-105 AFEX pretreatment, 

CBP process, varying 

pathways and coproducts

Gnansounou 

[17] 

2.12-

2.91

Straw, eucalyptus, 

poplar, 

switchgrass

1,760-

2,200

57-127 70-84 Dilute-acid 

pretreatment

Bals [18] 1.86-

2.20

Corn stover 850 45 78 AFEX pretreatment, 

varying pretreatment 

conditions

Piccolo [19] 3.43-

4.03

Hardwood 2,200 65 75 Dilute-acid 

pretreatment, varying 

financial inputs

KleinMarcuschamer 

[13] 

3.53-

4.58

Corn stover 1,700 60 52-74 Dilute-acid 

pretreatment, varying 

feed compositions 

and process 

conditions

National 

Academy of 

Sciences [20] 

1.20-

2.70 

(COP)

Poplar and “high 

glucan”

1,000-

1,600

50-88 67-106 Hot water 

pretreatment, Cost of 

Production analysis 

(COP<MESP)
10 

To further demonstrate the impact of such assumptions, we turn to the study of Kazi et al. [12], a 

joint effort between Iowa State University, ConocoPhillips, and NREL. While this paper 

examined a variety of pretreatment and downstream processing alternatives, its baseline was 

essentially NREL’s 2008 State of Technology (SOT) model (dilute-acid pretreatment, purchased 

cellulase enzyme, and C5/C6 co-fermentation) [21]. The Kazi et al. study adjusted several key 

parameters from the 2008 SOT model to incorporate external public data. The feedstock cost, 

enzyme cost, indirect capital cost factors, and reaction conversions were all modified per data 

external to NREL analysis. While both analyses were intended to represent near-term or “state of 

technology” economics extrapolated to an n

th plant as discussed previously, the assumptions in 

the Kazi et al. study were generally more conservative than NREL’s 2008 SOT, raising the 

MESP considerably. Adjusting for these key economic assumptions, their MESP closely

approaches that of the NREL 2008 SOT case, as shown in Figure 3. In many cases, economic 

variance between technoeconomic studies is easily explained by normalizing for a few important 

inputs. 

Figure 3. Adjusting the Kazi et al. base case ($3.40/gal) to NREL 2008 SOT assumptions 

1.6 About the NREL Aspen Model

The final version of the Aspen model used to generate this design report is named DW1102A; 

this version will be made publicly available along with this design report. We note that it is not 

uncommon to find subtle differences in Aspen results between software versions and patch 

levels. The results presented here were obtained with model DW1102A on Aspen Plus V7.2, 

patch level 0.

While Aspen can be completely rigorous, such detail is not always warranted in our simulation. 

Some unit operations, particularly solid-liquid separation, were modeled with a fixed 

$3.40 

$2.50 

$0.23 

$0.38 

$0.29 

$1.00 

$1.50 

$2.00 

$2.50 

$3.00 

$3.50 

$4.00 

Kazi et al. 

(published)

Indirect 

capital costs 

(adjust to 

2008 SOT)

Enzyme cost 

(70→32¢/gal)

Feedstock 

price ($75→ 

$60/ton)

Kazi et al. 

(adjusted)

MESP (2007$/gal)

2008 SOT = $2.43/gal
11 

performance determined by vendor testing. Bioreactors were modeled using experimentallydetermined conversions of specific reactions (e.g., cellulose to glucose) rather than using 

rigorous kinetics or rate expressions. This simple stoichiometric model still satisfies mass and 

energy balances.

The Aspen Plus simulation uses component physical properties internal to the software as well as 

property data developed at NREL or from the literature [22, 23]. Since the 2002 model was 

created, we have taken steps to reduce the number of custom-defined components and have 

eliminated the external property databanks. A discussion of the components and properties used 

is given in Appendix D. 
12 

2 Design Basis and Conventions

2.1 Plant Size 

The plant size in the present design is the same as in the 2002 design: 2,205 dry U.S. ton/day 

(2,000 metric tonne/day). With an expected 8,410 operating hours per year (96% uptime), the 

annual feedstock requirement is 773,000 dry U.S. ton/year. The 2002 study assumed that an 

ethanol plant would be responsible for collecting and storing its own corn stover supply for the 

year, and 2,205 dry ton/day represented a 50-mile corn stover collection radius, assuming that 

10% of the biomass within that radius was available to the plant. 

In the present design, feedstock is assumed to be delivered to the plant from a satellite storage 

and processing facility like the pioneer uniform-format facility described by Idaho National 

Laboratory (INL) [24]. That facility is designed to supply 882,000 dry U.S. ton/yr of biomass; 

enough to supply one of the ethanol plants described here or multiple smaller facilities. We have 

not repeated the plant-size sensitivity study from the 2002 report, which indicated that only 

modest cost reductions due to economies of scale could be achieved beyond 2,205 ton/day. 

Although the absolute MESP is different in the present report, we expect that the plant-size

trends observed in the 2002 report are still valid. 

2.2 Feedstock Composition

The feedstock type and composition can have significant impacts on the overall process design 

and economics. Feedstock type may influence the design of key components in the conversion 

process, e.g., the pretreatment reactor. Feedstock composition, from a potential sugar 

perspective, clearly affects the ethanol yield. The feedstock used for this analysis was corn 

stover. Stover is a generic term that refers to the “rest” of the corn plant, i.e., everything above 

the ground that is not a kernel of corn: leaves, stalks, husks, cobs, etc. NREL biochemical 

conversion research has focused on corn stover primarily because it is the most abundant 

agriculture residue in the U.S. and is thus readily available. While dedicated energy crops like 

switchgrass may ultimately have sustainability and yield advantages over agricultural residues 

like corn stover, these crops are currently not available in the large volumes required for pilotscale research. 

Corn stover can vary in composition and moisture content due to corn variety, region, weather, 

soil type, fertilization practices, harvesting and storage practices, time in storage, and so on. A 

recent NREL study assessed the compositions of corn stovers from many commercial hybrid 

corn varieties using a rapid compositional analysis method known as near-infrared

spectroscopy/projection-to-latent-structures modeling (NIR/PLS) [25]. This study evaluated the 

compositions of 508 commercial corn stover samples collected from 47 sites in eight Corn Belt 

states after the 2001, 2002, and 2003 harvests. From this study, the average concentrations (dry 

wt %) of the major stover components were 32% glucan, 19% xylan, 18% solubles, and 13% 

lignin (corrected for protein). These composition results are summarized in Table 3 along with 

the characteristic variation of each component. Figure 4 shows the distribution of total structural 

carbohydrate content among the samples. 
13 

Table 3. Summary of Whole Stover Composition Data 

Component Average 

(dry wt %)

Min Max Range

Ethanol solubles 3.3 1.7 4.1 2.4 

Sucrose 3.6 0 10 10 

Extractable inorganics (soil) 2.5 0 4.8 4.8 

Other water extractables 8.6 1.4 15.7 14.2 

Total solubles 17.9 5.7 30.8 25 

Glucan 31.9 26.5 37.6 11 

Xylan 18.9 14.8 22.7 7.9 

Galactan 1.5 0.8 1.9 1.1 

Arabinan 2.8 1.6 3.6 2 

Mannan 0.3 0 0.7 0.7 

Lignin (corrected for protein) 13.3 11.2 17.8 6.6 

Structural inorganics 3.9 0.8 6.6 5.8 

Protein 3.7 1.1 5.4 4.3 

Acetyl 2.2 0.9 2.9 2 

Estimated uronic acids 3.1 2.5 3.7 1.2 

Total structurals 81.6 70.4 90.8 20.4 

Component closure 99.5 93.8 104.9 11.1 

Figure 4. Structural carbohydrate distribution of the feedstock variability study [25] 

Table 4 compares the composition used in the 2002 design to the composition used in the present 

design. The composition used in the 2002 design was based on an average of nine stover samples 

obtained from two batches [26]. Compared with the data in Table 3, its glucan, xylan, and lignin 

concentrations were each higher than the more recent study’s averages for those components and 

near or above the maximum concentration measured in the 508 stover samples. (The earlier corn 

stover was handled extremely carefully between the field and the lab, while the samples from the 

more recent study were collected from the fields well after harvest.) The composition chosen for 

the present design was based on a single sample taken from the overall distribution. As indicated 

Total Structural Carbohydrates + Sucrose (dry wt %)

52 54 56 58 60 62 64

Frequency

0

10

20

30

40

50 Design 

Composition
14 

in Figure 4, this composition fits well inside the carbohydrate distribution of the more recent 

study and therefore provides a more representative design basis. The moisture content value 

assumed for corn stover was also updated to 20% based on the feedstock logistics specifications 

table developed by INL and reported in the OBP Multi-Year Program Plan (MYPP) [27]. 

Table 4. Corn Stover Composition from the 2002 Design [2] and the Present Design 

Component 2002 Design 

(dry wt %) 

Present Design 

(dry wt %) 

Glucan 37.40 35.05 

Xylan 21.07 19.53 

Lignin 17.99 15.76 

Ash 5.23 4.93 

Acetatea

 2.93 1.81 

Protein 3.10 3.10 

Extractives 4.68 14.65 

Arabinan 2.92 2.38 

Galactan 1.94 1.43 

Mannan 1.56 0.60 

Sucrose - 0.77 

Unknown soluble solidsb

 1.18 - 

Total structural carbohydrate 64.89 58.99 

Total structural carbohydrate + sucrose 64.89 59.76 

Moisture (bulk wt %) 15.0 20.0 a

Represents acetate groups present in the hemicellulose polymer; converted to acetic acid in pretreatment. 

b In the 2002 design, unknown soluble solids were calculated by difference to close the mass balance. This 

is now included in the extractives component. 

When converting the analytical composition to components used in the Aspen model, the waterand ethanol-soluble fractions from the compositional analysis were combined under 

“extractives.” The extractives component is assumed to be organic, with an average composition 

of CH2O, and consists primarily of sugars, sugar alcohols, and organic acids [28]. The presence 

of extractives in corn stover depends on the time of harvest and in part to how much microbial 

degradation of the material occurs after harvest; the amount of extractives in a given sample may 

therefore be indicative of its age. Additionally, where the mass balance did not sum to 100%, the 

extractives component was used to close it by difference. (In the 2002 design, the “unknown 

soluble solids” component performed this function.) Sucrose is another extractive component, 

but it is measured separately in laboratory analysis and has been added as a feedstock component 

in the present design; it is therefore excluded from the extractives component. The amount of 

sucrose present in corn stover is highly dependent on harvesting and handling practices. In 

pretreatment, this sucrose is assumed to be fully hydrolyzed to glucose and fructose. The fructose 

is further converted to degradation products in pretreatment, but the glucose resists degradation 

and thus is available for fermentation [29]. 

The differences in composition between the 2002 design and the present design have a few 

notable economic implications. Because the glucan (cellulose) content in the feedstock is lower, 

the present design has a lower ethanol yield per ton of feedstock. On the whole, the total 

structural carbohydrate composition (glucan + hemicellulose components) is about 5% lower 

than in the 2002 composition. Conversely, in the present design, a larger portion of the feedstock 

is non-carbohydrate organic compounds. These ultimately become process residues that are 
15 

burned to provide heat and electricity. The amount of electricity generated per ton of feedstock 

processed is therefore higher than in the 2002 design (though less electricity is exported as a coproduct because the present design has a higher internal electricity demand). 

2.3 Design Report Conventions 

2.3.1 Units 

The Aspen model we developed is, by legacy, based on the set of units required by Aspen for 

specifying custom component properties: kg, kmol, atm, °C for materials, and MMkcal (Gcal) 

for energies. Values in this report that were pulled directly from the Aspen model therefore tend 

to be reported in these units. Harris Group preferred to use U.S. standards (lb, Btu, °F, gal, etc.) 

when communicating with equipment vendors. Therefore, equipment specifications tend to be 

cited in these U.S. units. We have made an effort to cite both sets of units in this report, to the 

degree that it does not distract from the discussion. 

Note that in the present report, certain quantities (e.g., yields and costs) are computed and 

reported in terms of “tons.” To avoid ambiguity, tonne will denote a metric tonne (1,000 kg) and 

ton will denote a short or U.S. ton (2,000 lb). In general, the U.S. ton is the standard for this 

document. “Ton” also appears in Section 3.9 in the context of refrigeration, but this usage should 

be clear from the discussion. 

2.3.2 Total Solids Loading 

The process described here converts a solid feedstock (corn stover) into a liquid product 

(ethanol). Most material streams in the process therefore have a solid fraction and a liquid 

fraction. The relative amount of solids in a given stream is called its “solids loading.” Total 

solids loading is defined as the total weight percent of soluble solids (e.g., sugars and salts) and 

insoluble solids (e.g., cellulose and lignin) in a given material stream. Where useful, the total 

solids loading and the insoluble solids loading will be reported together. Note that in our 

convention, sulfuric acid, acetic acid, and ammonia are not considered soluble solids but 

ammonium acetate and ammonium sulfate are. Therefore, around some unit operations, e.g., 

hydrolysate conditioning, total solids loading is not a conserved quantity. 

2.3.3 Ethanol Density 

The most important results from this analysis are reported in terms of volume of ethanol 

produced: $/gal, gal/yr, gal/ton, etc. To avoid errors in these quantities that may arise from using 

the ethanol density computed by Aspen Plus, an accepted value was used for anhydrous ethanol 

at 20°C (68°F): 0.789 kg/L [30]. 

2.3.4 Theoretical Yields and Conversions 

The terms “yield” and “conversion” are used throughout this report to describe the extent of 

various chemical and biochemical reactions. It should be understood that each of these quantities 

is a percentage of the theoretical. For example, the statement “95% conversion of glucose to 

ethanol” means that 95% of the glucose in the system was consumed in the following reaction: 

Glucose  2 Ethanol + 2 CO2

which has a theoretical yield of 0.51 g ethanol / g glucose. 
16 

3 Process Design and Cost Estimation Details 

The process design described in this study was based upon demonstrated performance and 

planned performance goals from DOE’s core R&D efforts in biochemical conversion. This 

section describes the process in detail and discusses the influence of specific R&D goals in the 

decision-making process. 

3.1 Area 100: Feedstock Storage and Handling 

3.1.1 Overview 

Area 100 handles incoming biomass feedstock. In contrast to the 2002 design, which assumed 

bale delivery and yard storage of corn stover followed by milling and washing, the present 

design assumes that corn stover is delivered according to the specifications detailed in the Idaho 

National Laboratory (INL) design report for the pioneer uniform-format feedstock supply system 

[24]. In this envisioned design, biomass is stored in a central depot and is preprocessed and 

homogenized to a degree before delivery, such that the biorefinery receives feedstock with 

known, uniform-format specifications including particle size distribution, moisture content, and 

bulk density. 

Harris Group designed the receiving and handling system in Area 100 based on this delivery 

specification, using the INL design report as well as its own experience with the required 

equipment. The equipment in Area 100 is physically located at the ethanol plant and consists of 

weighing and unloading stations for incoming biomass supply trucks, short-term queuing 

storage, and conveyors for feeding bulk feedstock to the pretreatment reactor. A simplified 

process flow diagram for this area is shown in Figure 5, and more detail is given in PFD-P120-

A101 (Appendix E). 

Figure 5. Simplified flow diagram of the feedstock receiving system 

To pretreatment

Concrete storage dome

Dust collection

Truck tipper

Milled corn stover
17 

3.1.2 Design Basis 

In the uniform-format feedstock supply system design, feedstock would be stored in a satellite 

depot location with delivery to the biorefinery occurring six days a week by truck or possibly by 

rail. At the depot, material would be milled to a mean size of 0.16–0.23 in. (with a high content 

of fines), in order to achieve a mean bulk density of 9–11 lb/ft3 to maximize the biomass load per 

trailer. The incoming corn stover feedstock is assumed to have 20% moisture when it reaches the 

biorefinery; this is representative of a mixture of field-dried material having <15% moisture and 

co-harvested material having >20% moisture. Because the preprocessing operation is designed to 

lose very little dry matter and does not include any rinsing of biomass, it is assumed that the corn 

stover composition discussed in Section 2.2 is valid for the delivered material. We also note that 

the particle size of <0.25 in. is consistent with the material currently used in NREL pretreatment 

research. 

The as-received corn stover feed requirement for the plant is 2,756 U.S. ton/day (104,200 kg/h; 

229,700 lb/h) including moisture. In the projected design, refinery receiving operates on the 

same schedule as the biomass depot: 24 hours a day, six days a week. Each truck trailer holds 10 

U.S. tons of biomass. To satisfy production and storage requirements, the plant must receive 12 

trucks every hour. Incoming trucks are weighed by electronic scale (M-101) and unloaded using 

a whole-truck dumper (M-102) capable of a 7–10 minute unloading time. The dumpers empty 

into dedicated hoppers (M-103), which meter the biomass to a series of conveyors (C-101, C102, C-103). These carry material from the truck tipper to short-term storage. The minimum 

receiving rate is 244 ton/h to maintain 114 ton/h of continuous processing. Because trucks are 

not unloaded continuously, some extra capacity needs to be built into the conveyor that carries 

material from the truck dumpers to storage (minimum 330–440 ton/h recommended). In order to 

process 244 ton/h, and assuming a relatively constant flow of trucks, a pair of scales (one 

inbound and one outbound) and two truck dumpers are required. 

On-site storage is kept to a minimum of 72 hours to allow for a weekend buffer. Open piles are 

not favored due to concerns of fire, rodent infestation, and moisture degradation. Instead, the 

unloaded feedstock is stored in concrete domes (M-104). Two domes (each with a 36-hour 

capacity) are required so that one can be loaded while the other empties to the conversion 

process. 

Conveyors (C-106, C-107, C-108) connect the storage domes to the feedstock receiving bins on 

the pretreatment reactor in Area 200. A dust collection system (M-106) integrated with the 

conveyors and domes handles airborne particles released during the unloading and conveying 

processes. No dry matter is assumed lost in Area 100. In contrast to the 2002 design, the 

feedstock is not washed, which eliminates a source of evaporative water loss in this area. There 

is some evidence that feedstock washing may still be desired to reduce any incoming inorganics 

from fertilized feedstocks like corn stover, and this practice may be revisited in the future. 

3.1.3 Cost Estimation 

The feedstock cost assumed in this report is $58.50/dry ton (2007$). This cost comes from the 

Multi-Year Program Plan (MYPP) published by DOE’s Office of the Biomass Program [27]. The 

total of $58.50/ton includes $23.50/ton for a grower payment (MYPP Table B-1) and $35/ton for 

all collection, processing, storage, and transportation costs between the field and the receiving 
18 

bin on the pretreatment reactor (MYPP Table B-2). It should be stressed that these costs are 2012 

DOE research targets, like the conversion performance targets used in the Aspen model. 

Harris Group did obtain vendor quotes or other estimates for the truck unloading equipment, 

storage domes, conveyors, and dust collectors. Although the final equipment list assembled by 

Harris Group does not completely match the INL design report, the capital and operating costs 

for Area 100 are assumed to be included in the MYPP feedstock cost and are not included in the 

total capital investment (TCI) calculation; therefore they do not affect the MESP. 

The truck scales were quoted by St. Louis Scale, and the truck dumpers and hoppers were quoted 

by Jeffrey Rader. The storage domes, each with a 4,400 ton capacity, were quoted by Domtec. 

The internal reclaim system was quoted by Cambelt. All other connecting conveyors were 

quoted by Dearborn Midwest Conveyor Co. Sly, Inc. provided a quote for all components in the 

dust collection system. The belt scales for biomass leaving the storage domes were quoted by 

Tecweigh. 
19 

3.2 Area 200: Pretreatment and Conditioning 

3.2.1 Overview 

The pretreatment process converts most of the hemicellulose carbohydrates in the feedstock to 

soluble sugars (xylose, mannose, arabinose, and glucose) by hydrolysis reactions. Acetyl groups 

in the hemicellulose are liberated as acetic acid. The breakdown of biomass in pretreatment 

facilitates downstream enzymatic hydrolysis by disrupting cell wall structures, driving some 

lignin into solution, and reducing cellulose crystallinity and chain length. The nature and extent 

of such changes are highly dependent on the pretreatment chemistry and reaction severity 

(defined by residence time, temperature, and catalyst loading). Sugar degradation products such 

as furfural and 5-hydroxymethyl furfural (HMF) can also be formed in pretreatment. These 

compounds can have adverse effects on the fermenting organisms in sufficiently high 

concentrations. 

In the present design, hydrolysis reactions are catalyzed using dilute sulfuric acid and heat from 

steam. Pretreatment is carried out in two stages in the present design and the reaction severity is 

fairly mild compared to what has been modeled in the past. This milder pretreatment favors the 

production of soluble xylose oligomers. Higher-severity pretreatments tend to produce more 

monomeric xylose but at the risk of forming significant amounts of degradation products if 

reaction conditions are not well controlled. 

The first stage in pretreatment is a horizontal screw-feed reactor with a short residence time (5–

10 minutes). The second stage is a lower temperature/longer residence time “oligomer 

conversion” step that converts most of the xylose oligomers leaving the first stage to monomeric 

xylose without generating significant additional degradation products [31]. After the 

pretreatment reactors, the hydrolysate slurry is flash-cooled, vaporizing a large amount of water 

along with some of the acetic acid and furfural. The flash vapor is condensed and sent to the 

wastewater treatment area. The hydrolysate slurry is cooled by dilution water and sent to a 

conditioning reactor, where ammonia is used to raise its pH from 1 to 5–6. Figure 6 shows a 

simplified flow diagram of the pretreatment area. 

Figure 6. Simplified flow diagram of the pretreatment and conditioning process 

Milled corn stover

Acid

Steam

Steam

Dilution water

& NH3

Vertical presteamer

Horizontal reactor

Blowdown

tank

Oligomer

conversion

Ammonia

conditioning

To wastewater

To enzymatic hydrolysis

BFW

Hot water
20 

In previous designs, the pretreated biomass was separated into washed solid and liquid fractions, 

and the liquor and wash water were conditioned by an “overliming” process in which the pH was 

increased from 1 to ~10 with lime and then readjusted to ~5 with additional sulfuric acid. The 

lime and sulfuric acid precipitated as gypsum, which was removed in another solid-liquid 

separation step. The conditioned liquid was then re-slurried with the cellulosic solids before 

enzymatic hydrolysis. In overliming, a significant amount of sugar in the liquor (as much as 

13%) could be lost to side reactions occurring at high pH or pressed out with the wet gypsum. 

The present design uses ammonia in place of lime to avoid these sugar losses. The high 

miscibility of ammonia also permits conditioning of the whole hydrolysate slurry and eliminates 

the solid-liquid separation steps. Fermentation studies have indicated that there is no benefit to 

over-conditioning at high pH when using ammonia, so the hydrolysate is simply adjusted to 

enzymatic hydrolysis pH in one step. While ammonia is considerably more expensive than lime, 

the economic benefits of reduced sugar loss and reduced capital cost make ammonia the more 

economic alternative [32]. It is also possible that the ammonia could reduce nitrogen 

requirements in fermentation, but this currently is not modeled. 

3.2.2 Design Basis 

The pretreatment reactor system includes a feedstock receiving system, followed by a vertical 

vessel with a long residence time for steam-heating and potential acid impregnation of the 

biomass, followed by the horizontal pretreatment reactor, which operates at a higher pressure and 

a short residence time. PFD-P120-A201 (Appendix E) shows the components in greater detail 

and PFD-P120-A202 shows the oligomer conversion and conditioning equipment. The 

horizontal reactor configuration was chosen because it permits tighter residence time distribution 

control than a vertical reactor. This is important in single-step, high-severity pretreatment to 

minimize “over-cooking” or “under-cooking” portions of the biomass, either of which would 

lower the overall yield. 

Milled corn stover is fed to the receiving bins at the inlet of the pretreatment reactor (M-201). 

The receiving bins are rectangular with a bottom-drag conveyor that moves the feedstock pile 

inside the bin toward the doffing rolls (special scraper shafts with radial pins that spread material 

uniformly onto the discharge belt conveyor). The receiving bins also have a top rake-belt 

installed just under the bin cover to move incoming material toward the end of the bin, ensuring 

first-in/first-out material handling. At the discharge of each bin, the transfer conveyor (C-201) 

lifts material to the distribution screw conveyor (C-202), which supplies slightly (~10%) more 

feedstock than each pin drum feeder requires. Overfeeding allows the speed of the pin drum 

feeders to control the flow rate of material into each plug screw feeder. It also ensures that the 

plug screw feeders maintain a maximum fill. The excess feedstock is conveyed back to the 

storage bin by a series of screw conveyors (C-203). 

Each pin drum feeder (M-202) packs feedstock material into separate plug screw feeders (M203). The 26-in. plug screw feeder is a rugged, high-compression screw device designed to form 

a pressure-tight plug of material through axial compression. The compression action in each plug 

screw feeder may also squeeze out water and some of the water-soluble extractive components. 

This “pressate” could be handled as a separate, concentrated waste stream but in the present 

model, it is assumed that the pressate is recombined with the hydrolysate at the outlet of the 

reactor. Dilute sulfuric acid is injected at the discharge spool of the plug screw feeder. Feedstock 

drops from the plug screw discharge into a mixing and heating screw (C-204). The heating screw 
21 

discharges the feedstock into the top of the presteamer (M-204). Hot water is added at this point 

to control the pretreatment effluent at 30 wt % total solids. 

The vertical biomass presteamer is designed for a retention time of up to 10 minutes at a 

temperature of up to 165°C, though in the current model it only operates at 100°C such that no 

significant hydrolysis reactions occur in the presteamer. With a 2,205 ton/day throughput and 10-

minute retention time, a single vertical vessel is adequate. Feedstock flows downward through 

the vertical reactor with uniform temperature throughout. The reactor is discharged through a 

dual screw outlet device (C-205) to two plug screw feeders (M-206). The plug screw feeders 

meter feedstock to the pretreatment reactor and control the pressure and temperature difference 

between it and the presteamer. Acid for the pretreatment reaction is added at the discharge of 

each plug screw feeder. The transport conveyors (C-206, C-207) combine feedstock from both 

plug screw feeders and deliver it to the pretreatment reactor (M-207). 

The pretreatment reactor is a single horizontal reaction vessel. The reactor is designed for fairly 

severe conditions of up to 190°C (374°F) and about 1.1 wt % sulfuric acid. In the current design, 

the reaction conditions are milder at 158°C (316°F) and 18 mg acid/dry g of biomass. Acid is 

metered to the reaction chamber at a rate proportional to the mass flow rate of feedstock. Highpressure steam is injected into this vessel to maintain temperature. The reactor pressure is held 

just at the bubble point for the mixture. Heat losses from the reactor are not accounted for in the 

energy balance calculations. The residence time in the pretreatment reactor is nominally 5 

minutes. The reaction conditions are summarized in Table 5. 

Table 5. Pretreatment Reactor Conditions 

Sulfuric acid loading 18 mg/g dry biomassa

Residence time 5 minutes 

Temperature 158°C 

Pressure 5.5 atm (81 psia) 

Total solids loading 30 wt % a

 Additional acid is added downstream of the pretreatment reactor. 

The pretreatment reactor is discharged to a flash tank (T-203). The pressure of the flash is 

controlled to keep the temperature at 130°C (266°F). The slurry from T-203 goes into the 

secondary oligomer conversion reaction vessel (T-208), where it is held at 130°C for 20–30 

minutes. An additional 4.1 mg/g of sulfuric acid is added in the oligomer conversion step, 

bringing the total acid loading to 22.1 mg/g dry biomass. (Note that in the Aspen model, all of 

the acid is added to the pretreatment reactor block.) 

The oligomer conversion reactor is discharged into another flash tank (T-204) that operates at 

atmospheric pressure. After this flash, the hydrolysate whole slurry containing 30 wt % total 

solids and 16.6 wt % insoluble solids is sent to the conditioning tank (T-209). Here, the slurry is 

diluted with water to slightly greater than 20 wt % total solids to ensure miscibility through 

enzymatic hydrolysis. Ammonia gas is mixed into the dilution water to raise the hydrolysate pH 

to 5. The residence time in T-209 is 30 minutes and the dilution cools the slurry to 75°C (167°F). 

The flash vapor from T-203 is used to preheat the boiler feed water in H-812 or combined with 

the 100°C (212°F) flash vapor from T-204, which is condensed in H-201 and H-244. This 
22 

condensate contains volatile, potentially inhibitory organics created in pretreatment and therefore 

is routed to wastewater treatment (Area 600). 

Table 6 summarizes the reactions and percent conversions that take place in pretreatment. 

Glucan contained in the hemicellulose side-chains is converted to glucose along with a small 

portion of the cellulose. Minor hemicellulose carbohydrates (arabinan, mannan, galactan) are 

assumed to have the same reactions and conversions as xylan. The xylan-to-xylose conversion is 

an assumed total hydrolysis that also includes an enzymatic component that will be discussed 

later. The sucrose reaction to HMF and glucose reflects 100% hydrolysis of sucrose to fructose 

and glucose, followed by complete degradation of the fructose to HMF, as mentioned in Section 

2.2. 

Table 6. Pretreatment Hydrolysis Reactions and Assumed Conversions 

Reaction Reactant % Converted 

to Product

(Glucan)n + n H2O→ n Glucose Glucan 9.9% 

(Glucan)n + n H2O → n Glucose Oligomera Glucan 0.3% 

(Glucan)n → n HMF + 2n H2O Glucan 0.3% 

Sucrose → HMF + Glucose + 2 H2O Sucrose 100% 

(Xylan)n + n H2O→ n Xylose Xylan 90.0% 

(Xylan)n + m H2O → m Xylose Oligomera

 Xylan 2.4% 

(Xylan)n → n Furfural + 2n H2O Xylan 5.0% 

Acetate → Acetic Acid Acetate 100% 

(Lignin)n → n Soluble Lignin Lignin 5.0% a

 Sugar oligomers are considered soluble but not fermentable. 

3.2.3 Cost Estimation 

For the pretreatment reactor system, Harris Group obtained a detailed quote from Andritz, Inc., 

totaling approximately $20MM for the whole system: feedstock receiving bin, additional size 

reduction steps (if necessary), pre-steaming, pressurized heating, reaction, and flash cooling. A 

reactor schematic provided by Andritz is shown in Figure 7. This reactor configuration is similar 

in principle to the pretreatment reactor assumed in the 2002 design. 

The reactor system is constructed of carbon steel with all parts in contact with acid (pretreatment 

reactor, pressurized transporter, and plug screw feeder) clad in Incoloy 825. This met Andritz’s 

approval but was primarily chosen to be consistent with the 2002 design and a previous NREL 

corrosion study on reactor metallurgy [33]. This reactor system is actually overdesigned for the 

pretreatment process conditions considered here and is capable of operating at a significantly 

higher severity. In the present design, the pretreatment system contributes about $0.14/gal to the 

MESP; we chose not to adjust the cost of the reactor to account for the lower residence time, 

temperature, pressure, or acid loading. This will permit continued use of the same reactor cost in 

our analyses as researchers change the pretreatment severity in their optimization studies. As 

long as acid is used in any concentration, we believe it is unlikely that the reactor cost would 

change significantly because Incoloy (or similarly expensive) cladding would still be required. 

The reactor cost will be revisited in the future if pretreatment conditions change more 

significantly. 
23 

Figure 7. Horizontal pretreatment reactor design scheme (Andritz, Inc.) 

Costs for other equipment, such as pumps, agitators, tanks, conveyors, and heat exchangers, were 

estimated using Harris Group’s internal database. The holding tank for sulfuric acid (T-201) is 

plastic, and the rest of the equipment is stainless steel—most of it 316SS to provide additional 

acid resistance at temperatures below 100°C (212°F). 

3.2.4 Achieving the Design Case 

Table 7 shows demonstrated pilot-scale performance results achieved from 2008–2010 

[21, 34–35] as well as the 2012 conversion targets (i.e., those used in the Aspen model) for the 

pretreatment area. In pretreatment, the 2012 target is 90% conversion of xylan to monomeric 

xylose with 5% loss to degradation products. An additional 2012 target is to move to a wholeslurry conditioning process and thereby eliminate the need for a solid-liquid separation step in 

the pretreatment area. 

Feedstock

Feedstock

Steam

Flush water

Flush water

Steam

Acid

Vent

Hydrolysate

to oligomer 

conversion

Pressate

Pressate

reinjection

Screw feeder

Screw feeder

Vertical 

presteamer

Horizontal reactor
24 

Table 7. Research Status and 2012 Targets in the Pretreatment Area

2008

State of 

Technology

2009

State of 

Technology

2010

State of 

Technology

2012

Targets

Pretreatment

 Solids loading (wt %) 30% 30% 30% 30%

 Xylan conversion to xylose (%) 75% 84% 85% 90%

 Xylan conversion to furfural (%) 11% 6.4% 8% 5%

Conditioning

 Ammonia loading (g/L of hydrolysate) 13 10 4.8 4.8

 Hydrolysate solid-liquid separation yes yes yes no

 Xylose sugar loss (% entering conditioning) 2% 2% 2% 1%

 Glucose sugar loss (% entering conditioning) 1% 1% 1% 0%

In NREL’s pretreatment research, significant improvements in the conversion of xylan to xylose 

by dilute-acid pretreatment were demonstrated at the bench scale in 2007, where conversions of 

75% or higher were routinely achieved in bench-scale batch reactors [36]. Efforts to duplicate 

these results in continuous pilot-scale operation in 2008 proved more challenging. Continuous 

pretreatment is more difficult than batch pretreatment because process conditions (residence time 

in particular) cannot be as tightly controlled. Experiments in NREL’s 200 kg/day continuous 

horizontal pretreatment reactor indicated a xylan-to-xylose conversion limit of ~60% in that 

reactor configuration, though total conversion to monomeric and oligomeric xylose was around 

75%. The mild secondary oligomer conversion step was therefore added to convert the oligomers 

to monomeric xylose without generating additional degradation products, meeting the 2008 

xylose target of 75%. 

In 2009, the internals of the horizontal reactor were modified to tighten its residence time 

distribution [37]. Optimization of the operating conditions between the two stages met 2009 

targets with an overall xylan-to-xylose conversion of 79.6% with 6.4% loss to furfural. 

Remaining oligomers were 9.0%, leaving 5% of the xylan unreacted [38]. Further process 

optimization in 2010 showed little improvement over these conversions but additional xylose 

yield was obtained in enzymatic hydrolysis using an advanced cellulase preparation that has 

xylanase and “xylooligomerase” activities. In washed-solids enzymatic hydrolysis, the advanced 

enzyme preparation converted an additional 82% of the unreacted xylan to xylose. Overall, an 

85.3% yield of monomeric xylose was achieved from all thermochemical and enzymatic 

processes in 2010 experiments [39].

The targeted enzymatic conversion of xylan and xylose oligomers described above is quite 

promising [40]. It could eliminate the secondary oligomer conversion step used in our design 

and, in general, permit lower-severity pretreatments that yield more xylose oligomers and fewer 

degradation products. Lower-severity pretreatments have relaxed residence time control 

requirements and may be able to utilize lower-cost vertical reactor configurations. (Vertical 

pretreatment reactors are typically less expensive than horizontal reactors with the same 

throughput because they have less complex internals, though reduced severity will probably not 

eliminate the need for exotic reactor cladding on acid-wetted surfaces.) Additionally, lower 

pretreatment severity would result in lower energy requirements and reduced chemical costs, 

including reduced cleanup in wastewater treatment. 
25 

Including these hemicellulase “accessory” enzymes in our technoeconomic analysis is 

complicated because little to no cost information exists for these highly specialized, precommercial enzymes. However, preliminary analysis has shown that if the xylan and xylose 

oligomer hydrolysis is carried out simultaneously with the cellulose enzymatic hydrolysis, 

accessory enzymes are probably the more economic option, assuming they can be obtained at a 

cost per kg of protein similar to that currently assumed for cellulase [41]. 

Replacing overliming with ammonia conditioning has eliminated significant sugar losses and 

gypsum disposal cost. Because ammonia can be mixed with the dilution water and applied as an 

aqueous solution, the whole hydrolysate slurry may be treated at once without a solid-liquid 

separation step. However, from Table 7, one may note that elimination of hydrolysate solidliquid separation is not assumed until 2012. This is due not to problems with the conditioning 

procedure but to conversion limitations in whole-slurry enzymatic hydrolysis caused by endproduct inhibition, particularly with respect to the xylanase activities discussed above. 

Nevertheless, in the present design, which is intended to reflect 2012 technology, a whole-slurry 

process through conditioning and enzymatic hydrolysis is assumed. 

Further research topics in the pretreatment area include biomass deacetylation through acid preimpregnation, use of co-catalysts (e.g., iron salts), and disc-refining of the pretreated slurry. 

These are designed to reduce sulfuric acid usage in pretreatment and achieve high xylose yields 

under less severe pretreatment conditions while maintaining high enzymatic digestibility of 

cellulose in the pretreated solids, with economic benefits as discussed above. In addition, 

deacetylation and disc-refining may significantly enhance downstream enzymatic hydrolysis and 

potentially decrease enzyme loading requirements. Acid pre-impregnation is not explicitly 

modeled in our design, but it is envisioned that it could take place in the vertical presteamer, 

though the cost of the reactor system would have to be increased to include Incoloy cladding on 

the presteamer as well as all screw feeders and conveyors between the presteamer and the 

pretreatment reactor. This design effort is ongoing. 
26 

3.3 Area 300: Enzymatic Hydrolysis and Fermentation

3.3.1 Overview

In this process area, cellulose is converted to glucose using cellulase enzymes. This process is 

known as enzymatic saccharification or enzymatic hydrolysis. A cellulase enzyme preparation is 

a mixture of enzymes (catalytic proteins) that work together to break down cellulose fibers into 

cellobiose and soluble gluco-oligomers and ultimately into glucose monomers. The resulting 

glucose and other sugars hydrolyzed from hemicellulose during pretreatment are fermented to 

ethanol. Most fermentation research at NREL has used the recombinant co-fermenting bacterium

Zymomonas mobilis because this microorganism and much of its genome are in the public 

domain [42]. “Co-fermenting” means that the organism can simultaneously ferment glucose and 

xylose to ethanol. Other co-fermenting ethanologens are also being considered at NREL and 

elsewhere, e.g., metabolically engineered strains of Saccharomyces cerevisiae [43].

The process assumed in this design is known as separate (or sequential) hydrolysis and 

fermentation (SHF). Enzymatic hydrolysis is initiated while the slurry is still at an elevated 

temperature after pretreatment and conditioning. At this temperature, the enzyme activity is 

higher so conversion is faster and a smaller amount of enzyme is required. Once the conversion 

of cellulose to glucose is complete, the slurry is cooled to fermentation temperature and 

inoculated with the fermenting microorganism (the “ethanologen”). This deviates from the 2002 

process, which was based on a simultaneous saccharification and co-fermentation (SSCF) 

process in which the temperature of the slurry was reduced and fermentation was initiated before 

enzymatic hydrolysis was complete. In SSCF, the enzyme continues to hydrolyze cellulose even 

after fermentation is initiated. Fermentation creates a sugar sink (because sugar is being 

consumed by the ethanologen) that helps drive the enzymatic hydrolysis reactions toward 

glucose. 

The 2002 process was also nominally continuous with a total residence time of 3 days. Based on 

more recent experience with Z. mobilis, a 5-day SHF process with batch fermentation was 

selected as a more realistic case for the present design. Enzymatic hydrolysis begins in a 

continuous, high-solids reactor. As the cellulosic solids are saccharified by the enzyme, the 

viscosity of the mixture drops dramatically, such that it can be pumped to one of several parallel 

bioreactors. Hydrolysis continues in this vessel until complete, then the slurry is cooled and the 

ethanologen inoculum is added. The ethanol-containing fermentation broth is emptied to the beer 

well (storage tank) before being pumped to distillation. Figure 8 shows a simplified flow 

diagram of the enzymatic hydrolysis and fermentation process. 
27 

Figure 8. Simplified flow diagram of the enzymatic hydrolysis and fermentation process

3.3.2 Design Basis

The full process flow is shown in PFD-P120-A301 and -A302 (Appendix E). Enzymatic 

hydrolysis is initiated in a continuous, high-solids reactor (T-310). Neutralized, diluted 

hydrolysate is cooled with cooling water in H-301 and fed to T-310 at slightly higher than 

20 wt % total solids. A 100-hp inline mixer upstream of the continuous hydrolysis reactor 

(A-308) mixes the enzyme into the slurry. After the cellulase enzyme stream is mixed in, the 

total solids loading is 20 wt % (10.6 wt % insoluble) and the temperature is 48°C (118°F). The 

residence time in this first stage is 24 hours, after which the slurry is batched to one of twelve 

950,000-gal vessels (F-300), where enzymatic hydrolysis continues for another 60 hours. 

Moving pretreated material at 20% (or higher) total solids loading is challenging because this 

material is not pumpable until the cellulose has been partially enzymatically hydrolyzed. In the 

present design, the continuous high-solids hydrolysis reactor is envisioned as an empty tower, 

with the slurry entering at the top and flowing down with gravity. Horizontal mixers with 

rotating paddles are another design option for performing the initial enzymatic hydrolysis 

reaction, and NREL will be testing such a reactor in its new Integrated Biorefinery Research 

Facility (IBRF) in 2011 [44]. These are expected to be more expensive than the plug-flow tower 

specified here, but not dramatically so given the relatively short residence time expected in the 

high-solids vessels and when considered in the total equipment scope of the plant. We believe 

that the tower design is a suitable placeholder.

The amount of enzyme used (the enzyme loading) is determined by the amount of cellulose 

present in the hydrolysate and the specific activity of the enzyme. In the present design, the total 

cellulase loading is 20 mg enzyme protein/g cellulose to achieve a 90% conversion to glucose.

“Protein” here refers to the total concentration of protein in the enzyme broth as determined by 

assay; not all of this protein has cellulase activity. Although sugar yield generally increases with 

higher enzyme loading, the additional enzyme obviously comes at a cost. Enzyme loading 

therefore has a significant and complex impact on process economics. Determining the most 

economical enzyme loading requires optimization of several process parameters including 

temperature, residence time, and total solids loading [45]. The design conditions developed 

based on the recommendations of enzyme manufacturers and NREL researchers are summarized 

in Table 8. 

Beer

To recovery

High-solids hydrolysis

(continuous flow)

Batch Hydrolysis/Fermentation

Seed train

(Zymomonas mobilis)

Beer well

CW

Enzyme

Hydrolysate
28 

Table 8. Enzymatic Hydrolysis Conditions 

Temperature 48°C (118°F) 

Initial solids loading 20 wt % total solids (10.6% insoluble/9.4% soluble) 

Residence time 3.5 days total (84 h) 

Number and size of continuous vessels 8 @ 950 m3 (250,000 gal) each

Number and size of batch vessels 12 @ 3,600 m3 (950,000 gal) each

Cellulase loading 20 mg protein/g cellulose 

The batch reactors are agitated and temperature-controlled using a pump-around loop that 

consists of a centrifugal pump (P-300) and a heat exchanger (H-300). During enzymatic 

hydrolysis, the temperature is maintained with cooling water. Temperature control during 

enzymatic hydrolysis is important. Enzyme studies at NREL have indicated an optimum 

temperature of 48°C for a number of commercial and development enzymes. A -5°C deviation 

could result in a dramatically reduced rate of reaction; a +5°C deviation could result in the onset 

of protein denaturation. Either situation would reduce the cellulose conversion yield for the same 

residence time. 

Table 9 lists the reactions and conversions taking place during enzymatic hydrolysis. The 

saccharified slurry contains 11.7 wt % total soluble sugar (including oligomers), with 6.7% 

monomeric glucose and 3.7% monomeric xylose. As mentioned in Section 3.2, some of the more 

advanced enzyme preparations have shown a significant ability to also convert xylan to xylose. 

This additional conversion is captured in pretreatment so it is not included in Table 9. Note that 

glucose oligomers are modeled in Aspen as “glucose molecules in oligomeric form”; the actual 

size distribution of the oligomers is not captured.

Table 9. Enzymatic Hydrolysis Reactions and Assumed Conversions 

Reaction Reactant % Converted to 

Product

(Glucan)n → n Glucose Oligomer Glucan 4.0%

(Glucan)n + 1⁄2n H2O → 1⁄2n Cellobiose Glucan 1.2%

(Glucan)n + n H2O → n Glucose Glucan 90.0%

Cellobiose + H2O → 2 Glucose Cellobiose 100%

After 60 hours of additional enzymatic hydrolysis in F-300, the saccharified slurry is cooled by 

the pump-around loop and larger heat exchanger H-310 (specific to this task and shared between 

all the fermentors) to 32°C for fermentation. Recombinant Zymomonas mobilis bacterium is used 

as the ethanologen. This strain of Z. mobilis can simultaneously ferment glucose and xylose to 

ethanol. The present design assumes that the minor hemicellulosic sugar arabinose is also 

fermented to ethanol with the same yield as xylose, although this is a research target for 2012 

and has not yet been demonstrated at NREL. 

The inoculum protocol for Z. mobilis is a direct transfer of cells without a cell concentration step. 

In order to provide a required 10% inoculum volume back to the production fermentors, 10% of 

the saccharified slurry is split off to seed production (see PFD-P120-A301, Appendix E). Each 

seed train consists of five reactors operating in batch mode with a 24-hour batch time and an 

additional 12-hour turnaround time. The seed fermentors are cooled with chilled water from Area 

900 to maintain the temperature at 32°C (90°F). The first vessel (40 gallons) is inoculated with a 

seed culture from the lab. Its broth is used to inoculate a larger reactor, and so on. After five 
29 

iterations, the cell mass from the last vessel (200,000 gallons) is sufficient to inoculate the 

production vessel. Following the 2002 design, it was decided that two trains would be optimal

[2]. Using these conditions, the fifth reactor in Train A will complete fermentation at 180 hours. 

Train B will complete fermentation 12 hours later. Another 24 hours after that, Train A is again 

complete and the cycle continues. A discussion of seed train cycling can be found in the 2002 

report. Table 10 summarizes the seed train design specifications. Additionally, 0.1% (w/v) 

sorbitol is used in the final seed fermentor to improve cell viability at high sugar concentrations

[46, 47]. Sorbitol is not a component in the Aspen model, but it is reflected in the economics by 

purchasing an amount of sorbitol equivalent to 0.1% (w/v) of the sugar flow to the seed train, at a 

cost of approximately 0.7 cents/gal ethanol. 

Table 10. Seed Train Specifications

Inoculum level 10 vol % of production vessel size 

Batch time 24 h 

Fermentor turnaround time 12 h 

Number of trains 2

Number of fermentor stages 5

Maximum fermentor volume (F-305) 200,000 gal (757 m3) 

Corn steep liquor (CSL) loading 0.50 wt % 

Diammonium phosphate (DAP) loading 0.67 g/L fermentation broth (whole slurry)

Table 11 gives the reactions and conversions used in the seed fermentors to describe the 

microorganism growth and sugar metabolism. The fraction of sugar converted to cell mass is 

rather small. This is typical for Z. mobilis and is economically beneficial as well—most of the 

sugar entering the seed train is converted to ethanol, which is ultimately added to the production 

fermentors and recovered. So even though 10% of the saccharified slurry is diverted to produce 

the seed culture, its ethanol-producing potential is not lost. Other species (E. coli or yeast) could 

be expected to have cell mass yields about 2 times higher than that of Z. mobilis and thus 

produce modestly lower amounts of ethanol during seed culture propagation. 

Table 11. Seed Train Reactions and Assumed Conversions 

Reaction Reactant % Converted to 

Product

Glucose → 2 Ethanol + 2 CO2 Glucose 90.0%

Glucose + 0.047 CSLa + 0.018 DAP → 6 Z. mobilis + 2.4 H2O Glucose 4.0%

Glucose + 2 H2O→ 2 Glycerol + O2 Glucose 0.4%

Glucose + 2 CO2→ 2 Succinic Acid + O2 Glucose 0.6%

3 Xylose → 5 Ethanol + 5 CO2 Xylose 80.0%

Xylose + 0.039 CSL + 0.015 DAP → 5 Z. mobilis + 2 H2O Xylose 4.0%

3 Xylose + 5 H2O → 5 Glycerol + 2.5 O2 Xylose 0.3%

Xylose + H2O→ Xylitol + 9.5 O2 Xylose 4.6%

3 Xylose + 5 CO2 → 5 Succinic Acid + 2.5 O2 Xylose 0.9% a Corn steep liquor (CSL) and diammonium phosphate (DAP) are both nitrogen sources required for Z. mobilis

growth. The stoichiometry shown above is only used to balance the compositions assumed for nonstandard 

components like cell mass.

Two high-capacity transfer pumps (P-302) are used to transfer the seed to a holding tank (T-301) 

in a timely fashion (approximately 2.5 h). The seed hold tank (T-301) is sized to hold 20% more 
30 

than the fifth seed reactor (F-305). Another pump (P-301) feeds the seed to the production 

fermentation train. 

Besides being fermented to ethanol, sugar may be lost to side products by contaminating 

microorganisms. A total of 3% of the sugars available for fermentation is assumed lost to 

contamination, as discussed in the 2002 design report. This is the level of contamination loss 

tolerated by the corn ethanol industry, so it can be thought of as an expected value consistent 

with good batch fermentor operation. The contamination loss is modeled by converting a fraction 

of the sugars to lactic acid upstream of the fermentor. This allows us to simply assign a percent 

loss to contamination in the model so the conversions in the fermentor do not have to be 

adjusted. Table 12 shows the contamination reactions only. 

Table 12. Co-Fermentation Contamination Loss Reactions

Reaction Reactant % Converted to 

Product

Glucose → 2 Lactic Acid Glucose 3.0%

3 Xylose → 5 Lactic Acid Xylose 3.0%

3 Arabinose → 5 Lactic Acid Arabinose 3.0%

Galactose → 2 Lactic Acid Galactose 3.0%

Mannose → 2 Lactic Acid Mannose 3.0%

Fermentation is conducted in a batch system of 950,000-gal vessels (F-300). The fermentation 

residence time is modeled as 36 hours. Inoculum from the seed train is fed at 10 vol % of the 

hydrolysate flow along with the nutrients corn steep liquor (CSL, 0.25 wt %) and diammonium 

phosphate (DAP, 0.33 g/L of whole slurry). The fermentors are cooled using the pump-around 

loop with H-300 switched from cooling water to chilled water from the coolant loop (Area 900). 

Agitation requirements in the main production reactors were reduced from those in the 2002 

design. The main reactors (F-300) are agitated at 0.03 hp/1,000 gal (6 W/m3). The assumption is 

that most of the liquefaction and viscosity reduction of the slurry takes place in the continuous 

high-solids hydrolysis reactor, and vigorous agitation is not required in anaerobic Z. mobilis 

fermentation. 

Table 13 summarizes the conditions in the fermentation vessels and Table 14 lists the reactions

and conversions assumed in fermentation. Based on NREL’s knowledge of state-of-the-art cofermenting microorganisms, arabinose is assumed to have the same conversion to ethanol as 

xylose does. Conversion of galactose and mannose to ethanol is not assumed, in contrast to the 

2002 design. Arabinose follows the xylose stoichiometry, with 85% conversion to ethanol. 

Inhibition effects from furans and acetate are not modeled explicitly in Aspen, having no kinetics 

available to do so, but they are implicit in the static conversions used in the model because these

conversions are based on fermentation experiments carried out in the presence of these 

inhibitors. NREL does have an active research program in fermentation inhibitor characterization 

and mitigation [48], but the details of this program are beyond the scope of this process design 

report.
31 

Table 13. Co-Fermentation Conditions 

Organism Recombinant Zymomonas mobilis

Temperature 32°C (96°F)

Initial fermentation solids level 19.8% total solids (14.7% soluble, 5.1% insoluble w/w) 

Residence time 1.5 days (36 h) 

Inoculum level 10 vol %

Corn steep liquor (CSL) level 0.25 wt %

Diammonium phosphate (DAP) level 0.33 g/L fermentation broth (whole slurry) 

Table 14. Co-Fermentation Reactions and Assumed Conversions 

Reaction Reactant % Converted 

to Product

Glucose → 2 Ethanol + 2 CO2 Glucose 95.0%

Glucose + 0.047 CSLa + 0.018 DAP → 6 Z. mobilis + 2.4 H2O Glucose 2.0%

Glucose + 2 H2O → 2 Glycerol + O2 Glucose 0.4%

Glucose + 2 CO2 → 2 Succinic Acid + O2 Glucose 0.6%

3 Xylose → 5 Ethanol + 5 CO2 Xylose 85.0%

Xylose + 0.039 CSL + 0.015 DAP → 5 Z. mobilis + 2 H2O Xylose 1.9%

3 Xylose + 5 H2O → 5 Glycerol + 2.5 O2 Xylose 0.3%

Xylose + H2O → Xylitol + 0.5 O2 Xylose 4.6%

3 Xylose + 5 CO2 → 5 Succinic Acid + 2.5 O2 Xylose 0.9% a Corn steep liquor (CSL) and diammonium phosphate (DAP) are both nitrogen sources required for Z. mobilis

growth. The stoichiometry shown above is only used to balance the compositions assumed for Z. mobilis cell mass. 

Nutrient requirements have not been optimized and a minimal, low-cost nutrient formulation has yet to be defined.

The fermentation broth or “beer” has an ethanol concentration of 5.4 wt % and is collected in the 

beer well (T-306). A small amount of additional ethanol is entrained in the CO2 vent stream and 

will be recovered to the beer well via the vent scrubber. The beer well is designed for a residence 

time of 4 hours to provide surge capacity between fermentation and distillation. 

3.3.3 Cost Estimation

In general, the material of construction for all equipment in this section is 304SS, which is the 

most cost-effective material for fermentation service. Carbon steel has been proposed as an 

alternative and does have a precedent in the corn ethanol industry with some technology 

providers. However, stainless steel is more common, and Harris Group felt that carbon steel 

fermentors would have to be designed thicker for corrosion allowance, which would mostly

negate the capital savings from using the cheaper material. Carbon steel also tends to pit and can 

harbor biofilms, creating a contamination source. 

The continuous enzymatic hydrolysis reactor (T-310) is envisioned in this design as an empty 

tower that is loaded from the top. Harris Group obtained a quote from Caldwell for eight such 

vessels to be run in parallel. These were flat-bottomed plug-flow reactors with a 10:1 

height:diameter ratio. The larger vessels were quoted by Mueller: the million-gallon batch 

fermentors (F-300), the seed hold tank (T-301), and the beer well (T-306). Mueller also quoted 

the fourth and fifth seed fermentors (F-304, F-305), including internal cooling coils. The smaller 

seed fermentors (F-301, F-302, F-303) were quoted by A&B Process Systems as skidded units 

with cooling jackets and agitators. Harris Group developed costs for all of the pumps in this area 

using their internal equipment database. Costs for the larger agitators, i.e., those on the 

fermentors (A-300), the seed fermentors (A-304, A-305), and other large tanks (A-301, A-306), 
32 

were based on vendor (Lotus) quotes obtained by Harris Group. Large agitators tend to be topmounted because they are more energy-efficient than side-entry agitators. 

Harris Group obtained quotes from Alfa Laval for the plate and frame heat exchangers in this 

area (H-300, H-310). The incoming hydrolysate cooler is a shell-and-tube heat exchanger from 

Fox Engineering. 

3.3.4 Achieving the Design Case

For the enzymatic hydrolysis and fermentation area, Table 15 shows the demonstrated results 

used in the State of Technology (SOT) cases from 2008 to 2010 and the ultimate 2012 targets 

(i.e., those used in the Aspen model). 

Table 15. Research Status and 2012 Targets in Enzymatic Hydrolysis and Fermentation

2008

State of

Technology

2009

State of 

Technology

2010

State of

Technology

2012

Targets

Total solids loading (wt %) 20% 20% 17.5% 20%

Enzymatic hydrolysis substrate Washedsolids

WashedsolidsWashedsolidsWholeslurry

Combined hydrolysis and fermentation 

time (days) 7 5 5 5 

Corn steep liquor loading (wt %) 1% 1% 1% 0.25%

Overall cellulose to ethanol conversion 

(through hydrolysis and fermentation,%)a 85.5% 84% 87% 85.5%

Xylose to ethanol fermentation 

conversion (%) 80% 82% 79% 85%

Arabinose to ethanol fermentation 

conversion (%) 0% 51% 68% 85%

a

 Assumes 95% glucose-to-ethanol conversion. 

Enzymatic hydrolysis and fermentation experiments are typically “integrated” with pretreatment, 

meaning that a single large batch of pretreated material is used to perform many hydrolysis and 

fermentation experiments. Integrated performance results for the process configuration and 

operating conditions assumed in the 2002 design (i.e., 20% total solids loading with overliming) 

[2] were first produced in 2005. In 2008, integrated experimental results advanced significantly 

toward the 2012 targets following the replacement of overliming with ammonium hydroxide 

conditioning, which demonstrated a slightly improved cellulose-to-ethanol yield and 

significantly better xylose-to-ethanol yield due to decreased sugar losses and improved 

hydrolysate fermentability [32]. 

Enzymatic Hydrolysis

In all of the SOT experiments shown in Table 15, the pretreated hydrolysate underwent solidliquid separation and the liquor was conditioned with ammonium hydroxide as described 

previously. In these so-called washed-solids experiments, enzymatic hydrolysis was performed 

only on the pressed and washed pretreated solids (i.e., in the absence of pretreatment liquor). The 

conditioned liquor was added back to the hydrolysate slurry just before fermentation. The 2012 

targets for enzymatic hydrolysis, however, assume a whole-slurry process, in which the whole 

pretreated slurry is neutralized, hydrolyzed, and fermented without a complicated and costly

solid-liquid separation step. Based on bench-scale testing at NREL, Larox (now Outotec) helped 
33 

to estimate the purchased capital cost of a hydrolysate pressure filter system at the present design 

flow as being on the order of $30MM minimally. The solid-liquid separator cost used in the 2002 

design was only about $5MM. We suspect that the older cost estimate was extremely optimistic. 

In general, we have found a considerable discrepancy in the cost of solid-liquid separation 

technology being quoted now compared to what was quoted in the 1990s and 2000s, though we 

have not investigated why this might be so. 

In 2009, NREL began studying whole-slurry enzymatic hydrolysis in earnest because advanced 

enzyme preparations became available that could perform satisfactorily on this material. One 

advanced preparation was tested at the pilot scale and performed very well, achieving 91% 

cellulose-to-glucose conversion in initial process-relevant whole-slurry enzymatic hydrolysis 

experiments on a 2008 batch of pretreated corn stover (PCS) in which a 75% xylan-to-xylose 

conversion was achieved in pretreatment (see Table 7). Previously, cellulose-to-glucose 

conversions of 90% were not uncommon in washed-solids hydrolysis experiments, but this 

enzyme was the first tested at NREL to demonstrate such a high conversion in a whole-slurry 

experiment. 

As described in Section 3.2, the direction in NREL pretreatment research has been toward lowerseverity processes that rely on secondary thermochemical or enzymatic hydrolysis to convert 

xylose oligomers to monomers. The reduced severity also affects cellulose digestibility. In newer 

batches of PCS that have higher (≥80%) xylose yield, the glucose yield from cellulose is lower 

for the same total solids loading. Loss of conversion may be counteracted to a degree by 

reducing the solids loading or by employing pretreatment-related process steps such as 

deacetylation and/or disc-refining, as discussed in Section 3.2.4. Whole-slurry hydrolysis results 

using the advanced enzyme preparation are summarized in Table 16. 

Table 16. Summary of Integrated Whole-Slurry Enzymatic Hydrolysis Performance Using an

Advanced Enzyme Preparation 

SOT 

Year

PCS Batch Total Solids

Loading (wt %)

Glucose Yield

from Cellulose

2009 2008: 75% xylose 20% 91%

2009 2009: 80% xylose 20% 83%

2010 2010: 80% xylose 20% 84%

2010 2010: 80% xylose 17.5% 86%

2010 2010: 80% xylose 15% 90%

Cellulose-to-glucose conversion in these experiments is approaching the 2012 targets (90% 

conversion at 20 wt % total solids), but it is not there yet. Achieving these targets will require use 

of one of the rapidly-improving enzyme preparations being developed by NREL’s industrial 

collaborators or through DOE’s ongoing enzyme solicitation focused on the development of 

advanced enzyme systems [49]. In addition to cellulase activity, these advanced enzyme 

preparations will likely have hemicellulase activity, enabling less severe pretreatment. Accessory 

enzymes (e.g., ferulic acid esterase) may also be present in advanced cocktails to more 

efficiently hydrolyze the bonds within and linkages between cellulose, hemicellulose, and lignin. 

NREL will continue to monitor the performance of these evolving enzyme preparations as well 

as the cost of solid-liquid separation technologies in order to choose the most cost-effective 

process option for demonstration in 2012. 
34 

Fermentation

NREL continues to perform integrated fermentation experiments using the recombinant, cofermenting Z. mobilis strain 8b. As of 2010, the ethanol tolerance limit of this microorganism has 

become the major barrier in NREL fermentation research. In co-fermentation, this strain readily 

consumes all of the available glucose, and we assume 95% is converted to ethanol. This 

microorganism will further consume the xylose in the system, producing ethanol until it reaches 

its ethanol tolerance limit of approximately 60 g ethanol/L. (In practice, the presence of typical 

fermentation inhibitors from biomass hydrolysate drives the limit to <60 g/L.) 

As discussed above, sugar yields from pretreatment and enzymatic hydrolysis have been steadily 

increasing. This has led to lower percent xylose utilizations in fermentation experiments, 

especially those performed at higher total solids loadings where the xylose is more concentrated. 

Characteristic ethanol yields from xylose in the 2009 and 2010 SOT cases are shown in Table 

17. Even though the amount of available sugar changes in each case, the final ethanol 

concentration is more or less constant (within the accuracy of the measurement) and is indicative 

of the microorganism’s practical ethanol tolerance limit. At 20% total solids in 2009, the 

microorganism was able to convert 82% of available xylose before reaching its limit. In 2010, 

the same limit was reached after converting only 63% of available xylose. In other words, the 

results summarized in Table 17 suggest that Z. mobilis strain 8b was operating very close to its 

ethanol tolerance limit in 2009 and was most likely being inhibited by ethanol in 2010 when the 

sugar concentrations increased. 

Table 17. Ethanol Yields from Xylose

2009 SOT 2010 SOT

% Total Solids in 

Fermentation

Xylose 

Conversion

Final Ethanol 

Conc. (g/L)

Xylose 

Conversion

Final Ethanol 

Conc. (g/L)

15% - - 78% 49

17.5% - - 79% 53

20% 82% 52 63% 53

25% - - 20% 50

Ethanologen strain research is currently being performed outside of NREL through ongoing 

DOE solicitation projects [50]. The companies participating in these projects are tasked with 

improving the ethanol tolerance of co-fermenting microorganisms like Z. mobilis to alleviate the 

problem of reduced ethanol yield at high sugar concentrations. Despite the highly sensitive 

nature of strains (ease of propagation, value of intellectual property, etc.), NREL expects to 

receive some of these strains for testing in 2011 and we hope to be able to demonstrate better 

ethanol tolerance from one or more of these strains under integrated processing conditions. 

The fermentation strain solicitation awardees were also tasked with demonstrating improved 

conversion of the minor hemicellulosic sugars (arabinose, galactose, and mannose) to ethanol. 

Arabinose is the third most abundant sugar available in corn stover following glucose and 

xylose, making up about 5% of the total available sugar. Achieving an 85% arabinose-to-ethanol 

yield provides an additional 3 gallons of ethanol per ton of dry corn stover and decreases the 

MESP by about $0.08/gal compared to the MESP with zero conversion of arabinose. Galactose 

accounts for another 2.5% of the available sugar, and mannose content in corn stover is typically 

negligible. As of 2008, it was clear that the solicitation awardees were not likely to make 
35 

significant progress on galactose or mannose conversion but that conversion of arabinose to 

ethanol was possible through incorporation of known arabinose metabolic pathways. In 2010, 

one of the awardees reported to DOE that they had demonstrated arabinose conversions of 68% 

on pure arabinose and 90% when glucose was dosed along with arabinose. The specifics of these 

experiments are protected information and cannot be shared here. NREL has also begun a 

program to genetically engineer the arabinose fermentation pathway into the Z. mobilis 8b strain 

to reduce the risk that the proprietary ethanologens being developed through the DOE strain 

solicitation are unable to meet the 2012 cost targets [51].

Low-Cost Media

As seen in Table 15, corn steep liquor (CSL) is assumed as a nutrient source for ethanol 

fermentation. While NREL does not actually use CSL in routine laboratory or pilot plant 

experiments, its use is assumed in the technoeconomic model as a placeholder for a low-cost 

source of nitrogen and trace minerals, and reducing the quantity of this material used in 

fermentation is a 2012 target. CSL, a by-product of the corn wet milling process, is high in 

protein, vitamins, and minerals and is usually mixed with the corn fiber by-product and sold as 

gluten feed [52]. On its own, CSL is used as a nutrient source in many fermentation industries

including wet-mill corn ethanol, pharmaceuticals, and industrial enzyme production. Because 

CSL is a wet mill by-product, and the corn starch business is relatively static, there is in effect a 

fixed amount of CSL available in the United States. We have performed some projections 

elsewhere and have determined that it is unlikely that CSL could scale to support cellulosic 

ethanol production at commercial volumes of billions of gallons per year [51]. Finding another 

low-cost and scalable nutrient source is a subject of future NREL research and technoeconomic 

analysis.
36 

3.4 Area 400: Cellulase Enzyme Production

3.4.1 Overview

This process area produces cellulase enzyme that is used in Area 300 to hydrolyze cellulose into 

glucose. Cellulase refers to a mixture of enzymes (catalytic proteins) that includes: (1) 

endoglucanases, which attack randomly within the cellulose fiber, reducing polymer chain length 

rapidly; (2) exoglucanases, which attack the ends of highly crystalline cellulose fibers; and (3) βglucosidase, which hydrolyzes the small cellulose fragments (cellobiose, a glucose dimer) to 

glucose. Cellulase is produced industrially using (among other microorganisms) Trichoderma 

reesei, a filamentous fungus that secretes high levels of cellulase enzymes when grown 

aerobically in the presence of cellulose or other cellulase inducers. 

The 1999 design report [1] introduced Area 400 to produce cellulase on-site with ethanol, using a 

slipstream of the conditioned hydrolysate slurry from pretreatment. The 2002 design report [2] 

omitted Area 400 in favor of a purchased-enzyme model in which the enzyme was assigned a 

certain fixed cost contribution per gallon of ethanol. Area 400 was reinstated for this design 

report with new reactor design assumptions and several other improvements over the 1999 

cellulase enzyme production model. We note that by including an on-site enzyme production 

section, NREL and DOE are not making a judgment about whether or not the cellulosic ethanol 

industry should align to this mode of enzyme distribution. Rather, the model on-site enzyme 

section is intended to improve transparency in determining the true cost of cellulase enzymes for

large-scale cellulosic ethanol production. 

The present design considers submerged aerobic cultivation (“aerobic fermentation”) of a T. 

reesei-like fungus on a feedstock of glucose and fresh water. Producing cellulase enzymes with 

glucose is not straightforward, because the absence of cellulose does not encourage the 

microorganism to secrete cellulase enzymes. Here, we have assumed a media preparation step 

where a small fraction of glucose is converted to sophorose, a powerful inducer of cellulase, 

using a small amount of the cellulase enzyme itself. When grown on this substrate, T. reesei has 

been shown to productively secrete cellulase [53]. Using glucose as the substrate is not 

necessarily more expensive than using hydrolysate slurry from pretreatment because the 

expected enzyme titer is significantly higher with glucose, reducing the capital and utility costs 

compared with using hydrolysate. Figure 9 is a simplified flow diagram of the enzyme 

production section. 

Figure 9. Simplified flow diagram of the enzyme production process

Glucose + water

Sophorose conversion enzyme

To hydrolysis

Media prep

Aerobic culture

(fed-batch)

Holding tank

Glucose/sophorose Air

Seed train

(Trichoderma reesei)
37 

3.4.2 Design Basis

The on-site enzyme process was designed without input from enzyme companies or other protein

production experts, although several individuals knowledgeable about this topic did participate in 

the peer review of this report. Based on the claims in U.S. Patent 4,762,788 (Example 5) [54] and 

a few reasonable assumptions, we have developed a rudimentary process that is adequate for our 

needs. The key assumptions used in the current design are summarized in Table 18. Note that the 

mass yield of enzyme from glucose is actually a computed number that is dependent on 

assumptions regarding protein composition and microorganism-specific productivity, which will 

be described below.

Table 18. Area 400 Guiding Design Basis Assumptions 

Parameter Assumption

Protein loading to enzymatic hydrolysis 20 mg protein/g cellulose

Reactor size 300,000 L @ 80% final working volume

Enzyme titer at harvest 50 g/L

Mass yield of enzyme from glucose 0.24 kg enzyme/kg glucose

Enzyme production cycle time 120 h online, 48 h offline, 168 h total

The desired cellulase loading to enzymatic hydrolysis is 20 mg of enzyme protein per g of 

cellulose. “Protein” here refers to the total amount of high molecular weight protein in the 

enzyme broth as determined by assay; not all of this protein is active cellulase. The total protein 

demand was thus calculated to be 490 kg/h (1,080 lb/h). An additional 10% is produced to 

account for the slipstream provided to the media preparation tank to make the glucose/sophorose 

mixture. The size of the cellulase production vessels was chosen to be 300 m3

 (80,000 gal) with a 

height-to-diameter ratio (H/D) of 2. Fermentation is assumed to be a fed-batch process starting at 

50% working volume and ending at 80%. Over one week, each bioreactor will see a 24-hour cell 

growth period, a 96-hour protein production period, and a 48-hour offline period for draining, 

cleaning, and refilling. With a one-week total cycle time and the production parameters listed in 

Table 18, one bioreactor is capable of producing 12,000 kg of protein in a week, or 71.4 kg/h 

(157 lb/h). The equivalent enzyme volumetric productivity is 0.30 g protein/L-h. Nine reactors 

were therefore required to deliver the 490 kg/h of protein needed for enzymatic hydrolysis. 

The reactors are loaded initially with the glucose/sophorose carbon source and nutrients 

including corn steep liquor (CSL), ammonia, and SO2. After the initial cell growth period, 

additional substrate is added to maintain protein production. The bioreactors are sparged with 

compressed and cooled air and corn oil is added as an antifoam. The reactors are temperaturecontrolled by chilled water flowing through internal coils. 

To determine the aeration and agitation requirements for the production bioreactors, the oxygen 

transfer rate was set equal to the oxygen uptake rate (OTR=OUR), with the OUR computed 

stoichiometrically from the protein production rate. In the modeled bioreactors, the reaction 

stoichiometry balances the reactions of substrate, oxygen, ammonia, and SO2 to cell mass and 

enzyme (plus CO2 and water) using an elemental composition for commercial cellulase provided 

by Novozymes [55]. The composition of cell mass was taken as the average of a generic cell 

mass composition [56] and the enzyme composition, with the assumption that cell mass includes 

some unreleased protein.
38 

Cellulase protein: CH1.59N0.24O0.42S0.01

T. reesei cell mass: CH1.645O0.445N0.205S0.005

In the production bioreactors, it is assumed that 90% of the carbon source is converted via the 

protein reaction and 10% is converted via the cell mass reaction. In the seed reactors, 85% of the 

carbon is converted via the cell mass reaction and 5% via the protein reaction, with 10% 

unreacted. This represents an overall molar selectivity of glucose to 31% protein, 4% cell mass, 

and 65% CO2, yielding 0.24 kg enzyme protein/kg glucose. The total oxygen demand for one 

bioreactor producing 12,000 kg protein in 120 hours online is 8.4 kmol/h. At its maximum (i.e., 

at the initial working volume of 50%), the OUR is 56 mmol/L-h. 

The OTR is the product of the volumetric mass transfer coefficient kla (h-1) and the log mean 

oxygen concentration gradient (the difference between the saturated dissolved oxygen 

concentration and the desired dissolved oxygen concentration) at the top and bottom of the 

bioreactor operating volume: 

OTR ൌ ݇௟ܽ

௕ܥሺ

כ െ ܥ஽ሻ െ ሺܥ௧

ሻ஽ܥ െ כ

ln ൬

௕ܥ

஽ܥ െ כ

௧ܥ

஽ܥ െ כ

൰

where Cb* and Ct* are the saturated dissolved oxygen concentrations (mmol/L) at the bottom and 

top of the vessel, respectively, and CD is the desired dissolved oxygen concentration, which is 

assumed to be 30% of saturation at 1 atm absolute pressure (0.063 mmol/L). The top and bottom 

concentrations of saturated dissolved oxygen were calculated by Henry's Law: 

௅ܥܪ ൌ ߨ

where π is the partial pressure of oxygen (atm), H is Henry’s coefficient (atm/(mmol/L)), and CL

is the liquid concentration of the gas (mmol/L). Henry’s coefficient for oxygen in water was set 

equal to 1 atm/(mmol/L) for all calculations. (Henry’s coefficient for oxygen in pure water is 

approximately 1.3 atm/(mmol/L) at 25°C, but solubility may be affected by the non-water 

components in the broth; a coefficient of 1 is a reasonable initial estimate.) The pressure at the 

top of the vessel was taken to be 1 atm and the pressure at the bottom was determined by the 

static head of liquid in the vessel. The volumetric mass transfer coefficient was then computed 

using the following correlation [57]: 

݇௟ܽ ൫hିଵ൯ ൌ 8.42 ൬

ܲ

ܸ௅

൰

଴.ଷଷ

ܸீ

଴.ହ଺

where P is the impeller power consumption and VL is the working volume. The correlation 

assumes that P/VL is expressed in hp/1,000 L. VG is the superficial gas velocity (cm/s) in the head 

space. This correlation is intended for non-Newtonian fungal cultivation systems with a 

maximum volume of 30 m3

. Our bioreactors exceed the maximum intended volume for this 

correlation by several orders of magnitude but we believe that this equation yields results that are 

reasonable enough in the absence of more sophisticated reactor design involving computational 

fluid dynamics and so on. 
39 

Because we have assumed a fed-batch fermentation, the working volume varies over time, along 

with the OTR and aeration rate in vvm (standard volume of air per volume of liquid per minute), 

which are volumetric quantities. In absolute terms, the average aeration rate for a VG of 200 

cm/min was 0.83 standard m3/s (including two days of offline time). The average impeller power 

P was determined to be 260 kW (350 hp), with a maximum of 600 kW (800 hp). The final 

specifications for the cellulase production bioreactors are shown in Table 19. 

Table 19. Specifications of the Enzyme Production Bioreactors 

Total volume 300 m3 / 80,000 gal

Maximum working volume 80%

H/D ratio 2

Height 11.5 m

Diameter 5.75 m

Operating pressure 1 atm

Operating temperature 28°C (82°F)

Material 316SS

Agitator 800 hp

Four trains of three seed fermentors provide inoculum to the production bioreactors. Each vessel 

in the seed trains is run batchwise on the same substrate as the production vessels. Air is also 

sparged through each of the seed vessels, which are cooled with chilled water. The seed 

bioreactors are each sized at 10% of the next bioreactor volume, i.e., 0.3 m3, 3 m3, and 30 m3. 

The aeration demand is assumed to be 10% of the production aeration rate. Four trains were 

chosen because each production fermentor has a total cycle time of 7 days, while each seed 

fermentor should have a cycle time of 2 days (including cleaning and sterilization) to get through 

the cell growth phase only. 

Like the oxygen uptake rate, the glucose demand is also computed stoichiometrically from the 

required protein production rate. Ammonia and SO2 are fed to the reactors stoichiometrically and 

CSL, trace nutrients, and antifoam (corn oil) are added to the substrate based on flow rate. The 

required nutrient concentrations are taken from Schell et al. [58] and are restated in Table 20. 

Table 20. Cellulase Production Nutrient Requirements 

Component Amount

Ammonia Stoichiometric (approx. 7 g/L)

SO2 Stoichiometric (approx. 0.6 g/L)

CSL 1% w/w

Corn oil 0.1% v/v

Ammonium sulfate 1.4 g/L

Potassium phosphate 2.0 g/L

Magnesium sulfate 0.3 g/L

Calcium chloride 0.4 g/L

Tween 80 0.2 g/L

Glucose, the carbon source for cell mass and protein, is the most significant enzyme production 

expense in this model. Electricity is also a significant contributor due to the power requirements 

of air injection, agitation, and refrigeration. The electricity demand in the enzyme production 

area is summarized in Table 21. 
40 

Table 21. Cellulase Production Electricity Requirements 

Air compressor 1,408 kW

Agitators 2,326 kW

Chiller system (Area 400 contribution) 1,587 kW

Total electricity (includes pumps, etc.) 5,340 kW

Total electricity demand per kg protein 9 kWh/kg

3.4.3 Cost Estimation

Most equipment in this area is stainless steel. The air compressor and some of the nutrient 

delivery equipment items are specified as carbon steel. Harris Group obtained quotes for the 

production bioreactors (F-400) from Mueller, with internal cooling coils included. The 

production agitators and motors were quoted by Lotus. The smaller seed fermentors were quoted 

by A&B Process Systems as complete skid-mounted units with cooling jackets and agitators. The 

air compressor is a packaged unit from Dresser Roots, capable of delivering 8,000 SCFM of air 

at 3 atm. Harris Group developed costs for the pumps in this area using their internal database.

Not included in the enzyme production model are any costs for concentration, stabilization, or 

transportation of the enzyme to the plant. At the very least, one expects to have to pay licensing 

fees for the cellulase production microorganism, but these costs are not included because we lack

information on what they might be (though an educated guess might be 5 cents or less per gallon 

of ethanol). 

3.4.4 Enzyme Cost Discussion

The 2002 design model and subsequent analyses using that model assumed that cellulase 

enzymes had a fixed contribution to the MESP that was independent of the ethanol yield. This 

value was taken from the Multi-Year Program Plan (MYPP) of DOE’s Office of the Biomass 

Program [27], which in earlier editions stated that using technology available in 2009, the 

enzyme cost contribution should be $0.35/gal (2007$) and should decrease to $0.12/gal ethanol 

by 2012. These values were computed by NREL based on projections made by enzyme 

companies in 2005, following a DOE-funded research campaign that helped to significantly 

reduce cellulase enzyme production costs [59]. Enzyme costs in this range were considered 

realistic if production costs associated with transportation and formulation costs were eliminated, 

i.e., if the enzymes were produced on-site or near-site and were not formulated for stability 

during extended storage [60, 61].

Since the 2005 projections, DOE negotiated additional awards with several enzyme 

manufacturers [49] to further develop improved cellulase preparations and drive down 

production costs to the levels specified in the MYPP. In the process, DOE and NREL both 

received feedback from the enzyme industry that even $0.35/gal might be too optimistic for 

currently-available enzymes. Along those lines, in February 2010 the major enzyme 

manufacturers Genencor and Novozymes announced new commercial-grade cellulase enzyme 

preparations capable of higher performance at lower loadings. Both cited an enzyme cost 

contribution of approximately $0.50/gal ethanol [62, 63].

Based on the economics of the on-site enzyme section described above, the predicted cost of 

enzymes to the ethanol plant is $0.34/gal of ethanol. This is broken down as shown in Figure 10. 
41 

The carbon source (glucose) makes up 57% of the cost of the enzyme, i.e., $0.19/gal of the 

$0.34/gal is just for glucose. We can back out the unit cost of the enzyme from $0.34/gal by 

multiplying by the ethanol production rate and dividing by the enzyme flow rate, which works 

out to $4.24/kg protein. 

Figure 10. Enzyme production cost breakdown

The previous DOE target for enzyme cost was specified as $0.12/gal ethanol, or about one third

of what is reported here. One way to achieve such a low cost contribution would be to reduce the 

enzyme loading. In Figure 11, we see that the enzyme contribution to MESP is roughly linear 

with the enzyme loading (mg of enzyme per g of cellulose entering enzymatic hydrolysis). At the 

loading specified in the present design (20 mg/g), the enzyme contribution is $0.34/gal. The 

enzyme contribution approaches $0.12/gal at about 8 mg/g loading. Therefore, the original 

enzyme cost targets could potentially be achieved by a reduction in enzyme loading, but such a 

loading would be significantly lower than what we currently use to approach the 90% cellulose 

conversion yield target. 

Capital

21%

Glucose

57%

Other Nutrients

5%

Electricity

13%

Fixed Costs

4%

$0.34/gal ethanol

$4.24/kg protein

DW1102A
42 

 

Figure 11. Enzyme contribution to MESP as a function of loading

The lower limit of enzyme loading is not well understood. In theory, enzyme molecules are not 

consumed during the hydrolysis reaction but, over time, they may become inactivated by 

denaturation or irreversible binding to solid particles, especially when lignin is present. Benchand pilot-scale experiments at NREL are therefore typically overloaded (i.e., an excess of 

enzyme is used) to ensure complete hydrolysis, because we may only have the resources to run a 

limited number of experiments. However, we have observed that advanced enzymes demonstrate 

diminishing increases in conversion at loadings as low as 20 mg/g. Ultimately, the lower limit of 

enzyme loading will be determined by industry and enzyme manufacturers, and will likely 

decrease as pretreatment technologies and enzyme preparations continue to improve. For the 

present design, however, NREL was not comfortable modeling an enzyme loading less than 20 

mg/g. 

The enzyme cost contribution modeled here is lower than one would expect for an enzyme

preparation purchased from a separate, non-adjacent production facility. Transportation of the 

enzyme to the ethanol plant could alone add $0.09–$0.18/kg of product, even if formulation 

costs could be avoided. Furthermore, by lumping the enzyme production equipment in with the 

ethanol plant, some key items are inherently shared, e.g., the land and buildings, cooling tower, 

and utilities infrastructure. Overhead and fixed costs, especially labor and management, would 

also be higher for a standalone facility. Eliminating the shared aspects between the enzyme unit 

and the ethanol plant could easily add another 5 cents per gallon of ethanol to the enzyme 

contribution. Additionally, an external enzyme production facility would probably demand a 

higher rate of return than the 10% IRR assumed for the ethanol plant because it is a higher-risk 

and lower-volume business. Including all these extra costs would bring the total enzyme cost 

contribution in line with the $0.50/gal level cited in the press.

$-

$0.10 

$0.20 

$0.30 

$0.40 

$0.50 

$0.60 

$0.70 

0 10 20 30 40 50

Enzyme Contribution $/gal

Enzyme Loading mg/g cellulose
43 

Although the on-site enzyme cost contribution derived in our analysis is lower than what might 

be expected in a purchased-enzyme scenario, it does not necessarily make a strong argument for 

on-site production. The cellulase production unit described here, consisting of nine fermentors 

and four seed trains, is hardly a “plug-and-play” solution, and an ethanol producer could be 

forgiven for not wanting to assume the mantle of logistical and operational challenges associated 

with such a system. However, the cost benefits of co-location are real, and the economics of 

cellulase production are arguably so tightly coupled to those of the fuel product itself that an onsite cellulase facility might ultimately make economic sense in the future cellulosic ethanol 

industry. In the near term, however, it is more likely that ethanol plants will purchase enzyme 

from an external supplier who has an organization dedicated to improving enzyme performance 

and reducing costs. 
44 

3.5 Area 500: Product, Solids, and Water Recovery

3.5.1 Overview

Area 500 separates the fermentation broth from Area 300 into water, anhydrous ethanol, and 

combustible solids. Distillation and molecular sieve adsorption are used to recover ethanol from 

the raw fermentation beer and produce 99.5% ethanol. Distillation is accomplished in two 

columns. The first, called the beer column, removes the dissolved CO2 and most of the water. 

The second column is called the rectification column and concentrates the ethanol to a nearazeotropic composition. See Figure 12 for a process overview of this area. 

Figure 12. Simplified flow diagram of the separation process

The ethanol product leaves the rectification column at ~92.5% ethanol and is further dehydrated 

to 99.5% by vapor-phase molecular sieve adsorption. During regeneration of the adsorption 

columns, a low-purity (70 wt %) ethanol stream is generated and is recycled back to the 

rectification column for recovery. The beer column overhead stream is fed to a water scrubber 

along with the fermentation vents (containing mostly CO2 but also some ethanol), thereby 

recovering nearly all of the ethanol. The scrubber effluent is fed back to the beer well. 

The beer column bottoms (stillage) stream contains the unconverted insoluble and dissolved 

solids. The insoluble solids are dewatered by a pressure filter and sent to the combustor in Area 

800. The pressed stillage water cannot be recycled directly back to the process because it

contains high levels of organic salts such as ammonium acetate and ammonium sulfate, 

fermentation nutrients not consumed by the organism, and soluble inorganic compounds from 

the biomass. Instead, the pressed stillage water is directed to wastewater treatment (Area 600) for 

cleanup. 

ATM.

LP STEAM

CW

LP STEAM

CW

Fermentation Vents

Beer

Lignin to combustor

Stillage to WWT

Lignin Separator

Beer Column

Water to recycle

Regenerate

Ethanol product

Return to beer well

Clean water

Rectification 

Column Molecular Sieve 

Adsorption

Vent Scrubber
45 

3.5.2 Design Basis

Beer from the fermentation area (5.4 wt % ethanol) is fed to the beer column after passing 

through an economizing heat exchanger with the bottoms stream (PFD-P120-A501, Appendix 

E). The beer column (D-501) operates in a mode that removes the CO2 and as little ethanol as 

possible to the overhead while removing about 90% of the water to the bottoms. The ethanol is 

removed as a vapor side-draw from the column and fed directly to the rectification column (D502). Both columns are modeled with rigorous vapor-liquid equilibrium calculations in Aspen 

using the RADFRAC model. This is an improvement over the 2002 Aspen model, where the 

rectification column was modeled as a fixed separation block. This results in a more accurate 

determination of reboiler and condenser duties for the columns. We note that our Aspen model 

was not used by the distillation vendor (Megtec) in designing these columns—they arrived at a 

similar design using their own models. 

In the Aspen model, the beer column (D-501) contains 32 stages operating at 48% efficiency, 

with the feed entering on the fourth tray from the top. The molar reflux ratio required is 3:1. The 

overhead is vented to the scrubber and contains 85% CO2, 11% ethanol, and close to 4% water. 

All of the CO2 and only 0.3% of the ethanol entering the beer column is vented here. Most of the 

ethanol that is vented (99%) will be recovered in the vent scrubber and recycled. In addition, 

about 0.8% of the ethanol fed to the beer column is lost in the bottoms stream and is considered a 

permanent loss. More than 99% of the ethanol in the feed is removed as a vapor side-draw from 

the column at Tray 8 at 40 wt %. To minimize the ethanol loss in the bottoms, the beer column 

reboiler duty is kept relatively high, so there is a trade-off between ethanol loss and energy 

usage. Megtec designed the beer column with 32 fixed-valve trays. These have been found to 

tolerate solids well and have a reasonably good efficiency. The tray spacing is 24 in. and the 

column diameter is 14 ft. Heat is supplied to a forced-circulation reboiler (H-501) by lowpressure (9.5 atm [125 psig]) steam. The forced-circulation reboiler was considered necessary to

accommodate the solids present in the beer column bottoms. The beer column is operated at 2 

atm (15 psig) overhead pressure to keep the reboiler temperature low in order to minimize 

fouling. 

The vapor side-draw from D-501 is fed directly to the rectification column, D-502. This column 

uses 45 stages with an efficiency of 76%. The D-501 vapor is fed on Tray 33 (counting from the 

top). The recycle from the molecular sieve, which is higher in ethanol (72 wt % versus 40 wt %), 

is fed on Tray 14. The required molar reflux ratio is 3.5:1 to obtain a vapor overhead mixture of 

92.5% w/w ethanol and a bottoms composition of 0.05% w/w ethanol. Only 0.1% of the ethanol 

from fermentation is lost in the bottoms. The rectification column bottoms stream is recycled to 

the pretreatment reactor as dilution water; ethanol in this stream is not considered a loss. The 

7.5% water in the rectification column overhead represents only 0.7% of the water originally 

entering distillation. Megtec designed the rectification column with 45 valve trays. Above Tray 

15, the column is 13 ft in diameter; below Tray 15 it is 4.5 ft in diameter. (Structural supports for 

this tower are included in its installation cost.) Heat is supplied to the column in a thermosiphon 

reboiler (H-502) by low-pressure steam. For the present design, an air-cooled condenser was 

specified to condense the reflux. This was done to reduce evaporative losses from the cooling 

tower and will be discussed more in Section 5.2. 

The ethanol-containing vent from the beer column overhead, along with the vents from the seed 

and production fermentors in Area 300, is sent to a water scrubber (T-512). The scrubber column 
46 

recovers 99% of the vented ethanol. The scrubber effluent contains 1.8 wt % ethanol and is 

returned to the beer well. The scrubber is designed for absorption of >99% of the ethanol in the 

vents and no more than 40 tons per year of VOC (volatile organic compound) emissions for air 

pollution considerations. Higher water flow or lower water temperature could reduce VOC 

emissions from the vent scrubber, at a potentially significant increase in operating cost and water 

consumption. 

The overhead vapor from D-502 is fed to a molecular sieve adsorption unit (M-503), which is a 

system of columns packed with beds of adsorbent. Water is selectively adsorbed in the beds as 

the vapor flows through. The design specifics of this unit are not detailed here, but the general 

arrangement is given in PFD-P120-A503 (Appendix E). Saturated vapor from the rectification 

column is superheated and fed to one of two adsorption columns. The adsorption column 

removes 95% of the water and a small amount of ethanol. The 99.5% pure ethanol vapor is 

cooled by heat exchange with the regeneration condensate and finally condensed with cooling 

water and pumped to storage. While one bed is adsorbing water, the other is regenerating. The 

bed is regenerated by passing a small slipstream of pure ethanol vapor back through the watersaturated bed while a vacuum is applied. The ethanol strips the water off the adsorbent, and the 

mixture is condensed and returned to the rectification column (D-502). 

Finally, the beer column (D-501) bottoms stream is cooled to 47°C (117°F) in the feed/bottoms 

economizer; this is a suitable temperature for liquid-solid separation. The cooled material is 

separated in a pressure filter (S-505, PFD-P120-A505). The pressure filter is operated without a 

wash, and the cake is dried with air to 35% moisture. The drying air is sent to the combustor, and 

the pressed stillage water is sent to wastewater treatment in Area 600. This is a significant 

change from the 2002 process, in which the stillage was evaporated to a combustible syrup, with 

the evaporator condensate being recycled to the process. The effects of this change on the 

wastewater treatment area will be discussed in Section 3.6. 

A clean-in-place system (part of Area 900) services the equipment in this area that is prone to 

fouling (e.g., the beer column reboiler) by flushing with a hot caustic solution and circulating the 

solution until heat transfer surfaces are clean. Manual acid cleaning must be performed 

occasionally to remove mineral scale.

3.5.3 Cost Estimation

The construction material for equipment in the distillation area is generally 316SS. The beer 

column and rectification column were quoted by Megtec. The two columns were quoted as a 

package along with all pumps, heat exchangers (except the rectification column condenser), and 

reflux drums. The air-cooled rectification column condenser was quoted by GEA Rainey. The 

molecular sieve adsorption unit was quoted by Delta-T and the vent scrubber system was quoted 

by Envitech. The lignin separator was quoted by Larox as two FFP 2512 60/60 filter-press units. 

Harris Group used information and experience from other projects to estimate costs for the rest 

of the equipment, including solids handling equipment, pumps, and auxiliary tanks. 
47 

3.6 Area 600: Wastewater Treatment (WWT)

3.6.1 Overview

The ethanol process generates a number of wastewater streams that must be treated before 

recycle to the process or release to the environment. This is accomplished in the wastewater 

treatment system in Area 600. The treated water is assumed clean and fully reusable by the 

process, which reduces both the fresh makeup water requirement and discharge to the 

environment. 

Figure 13 is a simplified flow diagram of the WWT area. Condensed pretreatment flash vapor, 

boiler blowdown, cooling tower blowdown, and the pressed stillage water streams are mixed 

together. The combined wastewater stream is processed by anaerobic digestion and aerobic 

digestion to digest organic matter in the stream. Anaerobic digestion produces a biogas stream

that is rich in methane so it is fed to the combustor. Aerobic digestion produces a relatively clean 

water stream that can be reused in the process as well as a sludge that is primarily composed of 

cell mass. The sludge is also burned in the combustor. 

Figure 13. Simplified flow diagram of the wastewater treatment process

In previous designs, the stillage from the beer column was evaporated in Area 500 to yield a 

syrup that could be routed to the combustor, minimizing the loading to WWT in that model. In 

the present design, this process flow was ruled out because of the switch to ammonia 

conditioning in Area 200. In the 2002 design, sulfate ions introduced in dilute-acid pretreatment 

precipitated as gypsum after overliming and were removed from the process. With the switch to 

ammonia conditioning, the resultant ammonium sulfate and ammonium acetate salts remain in 

the process and end up in the stillage. When we modeled evaporation of the stillage to syrup, the 

salt concentrations in the syrup were so high that boiler vendors were not comfortable quoting a 

combustor system to handle it, due to concerns of burner fouling. We chose, therefore, to 

eliminate the stillage evaporators and route the stillage to wastewater treatment instead. 

While we were confident that anaerobic and aerobic digestion could satisfactorily treat the 

organic material flowing to WWT in our design, some technology vendors expressed concern 

with the concentrations of ammonia and sulfate and indicated that these could require special 

upstream cleanup steps, at additional cost and complexity. In 2010–2011, NREL therefore 

contracted with Harris Group and Brown and Caldwell (BC) to design a treatment system using 

actual material generated in the pilot plant. Their testing determined that upstream removal of 

Beer stillage

Cooling tower blowdown

Pretreatment flash

Boiler blowdown Biogas CH4/CO2

Air Vent to atmosphere

Digester sludge 

to combustor

Treated water

Aerobic

Digestion

Anaerobic

Digestion Clarifier

Dewatering

RO

Waste brine

Evaporator
48 

sulfur was not required for our process but that nitrification would be required in conjunction 

with aerobic digestion to handle the ammonia. 

The wastewater treatment technology described here deviates significantly from current corn 

ethanol plants, which strive toward zero-liquid-discharge designs [64]. Such plants may have 

aerobic digesters or even simple reverse-osmosis systems on internal recycle streams. The 

deviation in wastewater treatment technology here reflects the challenges inherent in cleaning up 

inorganics (fertilizer, etc.) that are found on the non-food portion of a food crop and corrosive 

chemicals added in the front end to overcome lignocellulosic biomass recalcitrance.

3.6.2 Design Basis

The wastewater treatment system was designed by Harris Group and Brown and Caldwell under 

the 2010 subcontract mentioned above. The design includes anaerobic digestion, aerobic 

digestion with nitrification, sludge dewatering, and reverse osmosis. Additional details can be 

found in the subcontract report [65]. The relevant design parameters of the wastewater system 

are summarized in Table 22. Some of these quantities were adjusted from what was determined 

by laboratory analysis to avoid artificially overdesigning the system for the rich fermentation 

media components and high suspended solids in the sample. 

Table 22. Wastewater Treatment System Design Basis

Hydraulic load 2.15 MMgal/d

Wastewater inlet temperature 50°C

Inlet pH 5.2

Total COD 87 g/L

Soluble COD 85 g/L

Total solids 68 g/L

Total suspended solids 1.5 g/L

Total alkalinity 2,750 mg/L as CaCO3

Ammonia-N 1,060 mg/L

Total Kjeldahl N 1,200 mg/L

Sulfate 4,400 mg/L

Silica 1,600 mg/L

Potassium 500 mg/L

Phosphate 805 mg/L

Upsets (including unusual suspended solids, fouling, etc.) and other intermittent loads (e.g., from 

process spills) were not considered in the design. Rain and snow run-off, equipment washing, 

and other non-process wastewater are assumed to flow to a municipal wastewater treatment 

system. Insoluble organic compounds like cellulose, xylan, and protein are assumed inert in 

anaerobic digestion, although these are largely removed upstream in the stillage lignin separator 

(Area 500). BC indicated that these solid components could be handled in WWT for roughly the 

same capital cost, but additional operating costs due to nitrification requirements of the protein 

fraction were prohibitive.

The process flow is detailed in PFD-A120-A601 and PFD-A120-A602 (Appendix E). The 

pressed stillage water, condensed pretreatment flash vapor, boiler blowdown, and cooling tower 

blowdown are mixed together and pumped through a heat exchanger (H-602) and cooled to 

35°C. The cooled stream then flows directly to the anaerobic digester (T-606). 
49 

The total chemical oxygen demand (COD) entering anaerobic digestion in the design case is 

approximately 26,500 kg/h (58,300 lb/h) (equivalent to 64 g/L). In anaerobic digestion, 91% of 

each organic component is destroyed; 86% is converted to biogas (methane and carbon dioxide) 

and 5% is converted to cell mass. Methane is produced from the utilized organic matter at 228 

g/kg COD removed. Carbon dioxide is also produced at a nearly equimolar rate to methane, such 

that the biogas from the digester is 51% CH4/49% CO2 on a dry molar basis. Cell mass is 

produced at a yield of 45 g cell mass per kg COD digested. Sulfate is converted to hydrogen 

sulfide (H2S) and leaves the system in the biogas. Since the COD-to-sulfate ratio in the design 

basis is ~20, the methane-producing organisms will out-compete the sulfate-reducing organisms, 

so upstream removal of sulfate is not necessary. 

During standard operation, the biogas will flow to the combustor, providing about 65 MMkcal/h 

(260 MMBtu/h) or 36% of the total fuel supplied to the combustor. The H2S in the biogas will be 

treated downstream of the combustor with flue gas desulfurization (Area 800). An emergency 

biogas flare is designed into the process to burn the biogas in case of a process upset in the 

combustor. 

The liquid from the anaerobic digester is pumped to aerobic activated-sludge lagoons with 

floating aerators (T-608). In the lagoons, 96% of the remaining soluble organic matter is 

removed, with 74% producing water and carbon dioxide and 22% forming cell mass. With a 

91% reduction of the COD in anaerobic digestion followed by a 96% reduction of the remaining 

COD in aerobic digestion, the total COD reduction is 99.6%. Ammonium ions from the stillage 

pass through to aerobic digestion where they are converted to nitrate by nitrifying bacteria. 

Because the nitrification process lowers the pH (by forming nitric acid), caustic must be added to 

the lagoons for neutralization. 

The fully digested material is pumped to a membrane bioreactor (R-609) for clarification. The 

membrane unit removes additional COD along with colloidal particles (especially silica). 

Coarse- and ultra-filtration upstream of the membrane unit separate the aerobic biomass sludge. 

Most of the sludge is recycled to the lagoons to maintain a high cell mass loading. The remaining 

portion is removed to a holding tank, where it is mixed with sludge from anaerobic digestion. 

The sludge in the holding tank is dewatered periodically using a centrifuge (S-611). The 

centrifuge solid is primarily cell mass and is conveyed to the combustor for disposal (C-614). 

The water leaving the centrifuge is recycled to the aerobic lagoons for additional treatment.

The treated water has a COD of about 0.5 g/L and is routed to a reverse osmosis (RO) membrane 

system for salt removal. Approximately 80,000 L/hr (350 gpm) of water is rejected from RO as a 

7% brine containing primarily sodium nitrate along with all remaining ions and organics. The 

RO water effluent is assumed to be pure and is mixed with process makeup water and recycled to

the process. The brine from RO is further concentrated in a mechanical-vapor-recompression 

evaporator (M-640) to 50% solids and the condensate is also recycled to the process. We do not 

further process the brine in our model at this time, and it is assumed to be waste. In reality, one 

may choose to further crystallize the brine for solid disposal. Some biomass burners may be 

designed to handle the salt (turning sodium nitrate into N2, NOx, and soda ash). It may also be 

possible to use flue gas to spray-dry this material. 
50 

3.6.3 Cost Estimation

Harris Group and Brown and Caldwell estimated cost for the wastewater system, drawing on 

support from technology vendors such as ADI and Biothane. While the BC estimate included a 

complete project cost work-up, we have only used the installed equipment costs in our 

technoeconomic model because we account for our own indirect and project costs at the plant 

level. To the extent that the cost data obtained by BC were publishable, details can be found in 

the subcontract report [65]. Harris Group also generated costs for the wastewater cooler (H-602) 

and sludge screw conveyor (C-614). 

3.6.4 Wastewater Challenges

The switch from overliming to ammonia conditioning in Area 200 has forced a major change in 

our WWT design. In overliming, the sulfate from sulfuric acid pretreatment was removed after 

conditioning in the form of “gypsum” (calcium sulfate). Previous work showed that this waste 

product was about 70% gypsum; the rest was unknown organics that could possibly be classified 

as hazardous waste requiring remediation (e.g., phenolics) [66]. The disposal cost of this material 

was therefore not well defined in the 2002 model, and ultimately the sugar losses associated with 

overliming were unacceptable, prompting the switch to ammonia conditioning. 

Because the stillage containing ammonium and sulfate salts could no longer be concentrated into 

a combustible syrup, as noted earlier, the chosen alternative was to route the stillage to 

wastewater treatment. The ammonia is removed from the system as sodium nitrate (which 

requires caustic) and the sulfate is removed as gypsum from the combustor flue gas (which 

requires lime). At this point, readers may question the wisdom of replacing lime in pretreatment 

with the combined requirement of ammonia in pretreatment, caustic in WWT, and lime in FGD; 

this is a valid and open question. However, NREL feels that going back to overliming is not the 

obvious solution, given the cost of hydrolysate solid-liquid separation, uncertainty around the 

cost of gypsum disposal, and staggering sugar losses associated with overliming. We will 

investigate how these options affect process economics and sustainability in a future study. The 

primary research path, as discussed in Sections 3.2 and 3.3, is to reduce the amount of chemicals 

added in pretreatment and thereby move to so-called no-conditioning processes that may be 

applicable given the advanced enzymes and fermentation organisms currently in development. 

This should reduce or eliminate the need for removal of inorganic material from wastewater and 

bring the plant more in line with current corn ethanol technology. 
51 

3.7 Area 700: Product and Feed Chemical Storage

3.7.1 Overview

This portion of the plant provides bulk storage for process chemicals and the ethanol product. 

The chemicals stored in this area include ammonia, corn steep liquor (CSL), sulfuric acid, and 

gasoline (used as a denaturant for the ethanol product). Water for fire suppression is also stored 

here.

3.7.2 Design Basis

Table 23 shows the major storage requirements for the present design. PFDs for the storage area 

can be found in PFD-P120-A701 and PFD-P120-A702 (Appendix E). 

Table 23. Storage Requirements 

Material Tank No. Size

Ethanol product T-701 Sufficient to contain 7 days of production: 2 carbon steel tanks @ 

750,000 gal each

Gasoline denaturant T-710 7 days of blendstock: 1 carbon steel tank @ 65,000 gal

Sulfuric acid (93%) T-703 5 days of storage: 1 carbon steel tank @ 12,600 gal

Ammonia T-706 5 days of storage: 2 SA-516-70 tanks @ 28,000 gal. Ammonia is 

stored anhydrous at 250 psig.

Fire water T-704 4 hours of fire suppression @ 2,500 gpm: 1 glass-lined carbon steel 

tank @ 600,000 gal

Corn steep liquor T-720 5 days of storage: 1 glass-lined carbon steel tank @ 70,000 gal

Diammonium 

phosphate (DAP)

T-760 7 days of storage: 1 SS304 tank @ 12,800 gal

DAP is assumed to be packaged in super-sacks and is unloaded via solid feeder (C-755), 

requiring an unloading blower (P-755) and DAP vent baghouse (S-755). From the storage bin, 

DAP is fed into a day tank (T-760) where it is mixed with water with a 5-hp agitator. The DAP 

solution is pumped (P-760) to fermentation. The ethanol product pump (P-701), sulfuric acid 

pump (P-703), ammonia pump (P-706), CSL pump (P-720), and gasoline pump (P-710) are sized 

for quick loading and unloading of trucks. The fire water pump (P-704) is sized for 2,500 gpm. 

3.7.3 Cost Estimation

Harris Group developed costs for most of the equipment in this section using their internal cost 

database. Quotes for some of the more specialized-material tanks were obtained from Mueller or 

Trinity Tank (via Chemithon). Quotes for agitators for the DAP and CSL tanks were obtained 

from Lotus, Inc. 
52 

3.8 Area 800: Combustor, Boiler, and Turbogenerator

3.8.1 Overview

The purpose of the combustor, boiler, and turbogenerator subsystem is to burn various organic 

by-product streams to produce steam and electricity. Combustible by-products include all of the 

lignin and the unconverted cellulose and hemicellulose from the feedstock, biogas from 

anaerobic digestion, and biomass sludge from WWT. Burning these by-product streams to 

generate steam and electricity allows the plant to be self-sufficient in energy (“thermal-neutral”), 

reduces solid waste disposal costs, and generates additional revenue through sale of excess 

electricity.

The fuel streams are fed to a combustor capable of handling the wet solids. A fan moves air into 

the combustion chamber. Treated water enters the heat exchanger circuit in the combustor and is 

boiled and superheated to high-pressure steam. A multistage turbine and generator are used to 

generate electricity. Steam is extracted from the turbine at two different conditions for use in the 

process. In the final stage of the turbine, the remaining steam is taken down to a vacuum and 

condensed with cooling water for maximum energy conversion. The condensate is returned to 

the boiler feed water system along with condensate from the various heat exchangers in the 

process. The steam turbine turns a generator that produces AC electricity for all users in the 

plant. The balance of electricity is assumed to be sold back to the grid, providing a co-product 

credit.

3.8.2 Design Basis

The moisture of the combined solid feed to the combustor is 44%; a small amount of additional 

moisture is entrained in the biogas. Given the fuel disposition, Babcock & Wilcox recommended 

a TowerPak Stirling Power Boiler system. This system features a live-bottom grated fuel bin to 

ensure drying and complete combustion of the wet solid fuel. The system was quoted to 

produced 525,000 lb/h (239,000 kg/h) of steam at 850°F (454°C) and 900 psig. Boiler efficiency, 

defined as the percentage of the feed heating value that is converted to steam heat, is ~80%. 

Flue gas from the combustor preheats the entering combustion air then enters a spray dryer for 

flue gas desulfurization (FGD). Due to the switch to ammonia conditioning, the sulfur content in 

the combustor feed is significantly higher than it was in the 2002 design. All of the sulfur 

entering the combustor is converted to sulfur dioxide, and its concentration in the flue is >1,800 

ppmw. This level of SO2 requires FGD. Lime (calcium hydroxide) is sprayed into the flue gas as 

a 20 wt % slurry at 20% stoichiometric excess. FGD converts 92% of the SO2 into calcium 

sulfate, which falls out the bottom of the spray dryer. The water in the slurry is vaporized and 

exits with the flue gas to the baghouse. The baghouse removes particulate ash, which is 

landfilled along with the calcium sulfate. The scrubbed gas is exhausted through a stack. 

Pretreatment research is currently targeting a lower acid use, so we will revisit the issue of SO2

emissions in the future. Additionally, 1% of the generated sulfur dioxide is converted to sulfuric 

acid. The presence of this acid requires that the flue gas temperature be kept above the dew point 

to avoid corrosion. 

Other pollutant generation rates were not changed from the 2002 design [67] and are computed 

based on the heating value of the combustion feed stream. Carbon monoxide is assumed to be 
53 

generated at a rate of 0.31 kg/MWh (0.2 lb/MMBtu). Nitrogen oxide, or NOx, is generated at 

0.31 kg/MWh (0.2 lb/MMBtu). NOx formation is a complicated mechanism and depends on the 

feed, combustion temperatures, excess air rate, combustor design, and the presence of flue gas 

cleanup devices like FGD. The fate of the ammoniated salts in the combustor feed is not well 

understood, but it is likely that some of the ammonia is converted to nitrous oxide (N2O) while 

also acting to reduce NOx formation [68]. (Ammonia injection is indeed one way to reduce NOx

emissions.) 

On the steam side, the boiler feed water (BFW) system includes a softener (M-820) for the 

makeup and condensate water, a deaerator (T-826) to remove air and other non-condensables, 

surge tanks, and pumps. Makeup water and condensate are treated and deaerated in these units 

then preheated with steam to 177°C (351°F). The boiler feedwater pumps (P-826) raise the BFW 

pressure to 900 psig (62 atm), requiring 540 kW (720 hp) of power. An amine injection system is 

specified to remove oxygen, which can pit the boiler surfaces. Other boiler chemicals include 

ammonia, added to control the pH and reduce the corrosive nature of the hot condensate, and 

phosphate, added to control the formation of scale in the steam drum. Boiler blowdown is 3% of 

steam production. The blowdown stream is flashed to atmospheric pressure; the vapor is vented 

to the air and the liquid is routed to wastewater treatment.

The turbogenerator (M-811) uses a multistage turbine with two extraction ports and a final 

condenser. Of the superheated steam leaving the boiler, 12% is extracted from the turbine at 175 

psig (13 atm) and 268°C (514°F) for feeding to the pretreatment reactor and the BFW 

economizer. An additional 35% is extracted at 125 psig (9.5 atm) and 164°C (327°F) and is used 

in distillation and in the deaerator. The rest of the steam is condensed at -13 psig (0.1 atm). The 

condensate is pumped back to the boiler. The turbine shaft turns a generator to produce 

electricity. The generator efficiency is assumed to be 85%. For this design, a total of 41 MW of 

power is generated from the system. The process uses 28 MW, leaving about 13 MW to be sold 

to the grid. 

3.8.3 Cost Estimation

Harris Group investigated the combined system of combustor, boiler, and turbogenerator. They 

contacted several vendors to obtain cost and operating information to identify a system suitable 

for the available fuel streams and steam requirements. The combustor/boiler quote ultimately 

came from Babcock & Wilcox and included the BFW preheater, FGD spray dryer, and 

baghouse. For the baghouse, bag replacement appears as a periodic charge in the cash flow 

worksheet. 

The turbogenerator is essentially an off-the-shelf technology, and Harris Group obtained a quote 

for a 42.2 MW Siemens generator suitable for this project. Harris Group also obtained quotes 

from Proctor Sales for support equipment including the deaerator, chemical injection system, 

tanks, and pumps. 
54 

3.9 Area 900: Utilities

3.9.1 Overview

Area 900 is used to track the utilities required by the ethanol production facility (except for

steam, which is provided by Area 800). Area 900 tracks cooling water, chilled water, plant and 

instrument air, process water, and the clean-in-place (CIP) system. Area 900 also tracks the 

electricity usage throughout the plant. 

The process water manifold in Area 900 mixes fresh water with treated wastewater (assumed 

suitable for all plant users) and provides this water at a constant pressure to the facility. Water is 

provided to the cellulase production unit, boiler and cooling tower makeup, the CIP system, and 

the vent scrubber. Fresh water is also mixed with some internally-recycled water for dilution 

before pretreatment and enzymatic hydrolysis. The plant and instrument air systems provide 

compressed air for general use (pneumatic tools and clean-up) and instrument operation. Larger 

users of compressed air, namely the stillage filter press and cellulase system, have their own 

compressors specified. The CIP system provides hot cleaning and sterilization chemicals to 

hydrolysis and fermentation, the enzyme production section, and the distillation system. 

3.9.2 Design Basis 

The cooling water system is designed for a 28°C supply temperature with a 9°C temperature rise 

in coolers throughout the facility. This is an assumed average rise; the actual cooling water rises 

across each exchanger are not explicitly modeled in Aspen. The primary cooling water users in 

this process are listed in Table 24. The percentage of cooling duty contributed by each user is 

shown in Figure 14. 

Table 24. Cooling Water Users

M-811 Condenses the steam turbine exhaust at a vacuum.

M-908 The chilled-water loop requires cooling water to condense the refrigerant. The cooling 

water duty to M-908 is set equal to the total load on the chilled-water loop.

H-301 Cools the pretreated slurry to enzymatic hydrolysis temperature (48°C).

H-509 Cools the ethanol product leaving the molecular sieve before it enters storage.

F-300 Before fermentation is initiated, the slurry must be cooled to fermentation temperature. 

It is assumed that cooling water would be used for the initial drop from 48°C to 35°C 

and that chilled water would afterward be substituted to sustain this temperature.

H-503 Provides cooling and condensation during regeneration of the molecular sieve.

M-401 Uses cooling water to cool the compressed air used in enzyme production.

H-504 Condenses the beer column reflux.

H-244 Condenses residual pretreatment flash vapor before it enters WWT. Most of the vapor 

is condensed by heat exchange with the incoming dilution water and boiler feed water.
55 

Figure 14. Cooling duty distribution between major users

By far the biggest user of cooling water is the condensing turbine. In contrast to the 2002 design, 

the cooling water duty required by the ethanol separation section is relatively small because an 

air-cooled condenser was specified for the rectification column. Aspen computes the cooling 

tower evaporation rate based on a temperature drop from 37°C to 28°C. It was assumed that 

windage would be 0.005% of the total flow to the tower. The tower blowdown was assumed to 

be 0.15% of the flow leaving the tower basin. 

Chilled water is provided by two 2,350-ton Trane centrifugal chillers. Per the chiller spec sheet, 

the compressor electricity demand for the chiller was estimated at 0.56 kW/ton of refrigeration. 

The cooling water demand for the chiller system was assumed to be equal to the heat removed in 

the chilled-water loop. The chilled-water loop provides cooling to the fermentors in Area 300 

and Area 400. 

Fresh water is assumed to enter the facility at 13°C and is used to cool the wastewater entering 

Area 600 to digestion temperature before entering the process water tank. The fresh water is 

mixed with the treated wastewater effluent in the process water tank (T-914) and then split 

several ways. Clean water must be provided to biomass dilution in the front end, to the cellulase 

production unit, to the vent scrubber, and to the boiler and cooling tower water makeup. The 

process water tank is designed for an 8-hour residence time. The process water pump (P-914) 

pumps water from the tank into the facility and is designed to handle 1.5x the process water flow

requirement. In the present design, less water is recycled internally in the process than in the 

Condensing 

Turbine

52%

Chiller condenser

14%

Hydrolysis 

cooler

11%

Hydrolyzate flash 

condenser

8%

Pre-fermentation 

cooler

6%

EtOH product 

cooler

5%

Molsieve 

regeneration 

cooler

3%

Enzyme production 

compressor

1%

Beer column 

condenser

0%

DW1102A

Total Duty = 97 MMkcal/hr (386 MMBtu/hr)
56 

2002 design. The only water stream directly recycled to the front end of the process is the 

rectification column bottoms.

The plant and instrument air systems provide compressed air for pneumatic tools and clean-up 

and instrument operation. The plant air compressor is sized for 400 SCFM at 125 psig. An 

instrument air dryer and surge tank were designed to provide clean dry air at a consistent 

pressure to the instrument air system. The surge tank was sized at 3,800 gal. 

About 68% of the electricity generated by the boiler in Area 800 is used throughout the plant to 

power pumps, agitators, compressors, etc. The surplus is sold to the grid for credit. The 

distribution of total plant power utilization among all areas is shown in Figure 15. Note that the 

cost of the power required by Area 100 is already assumed to be included in the feedstock cost 

but must be subtracted from the plant’s electricity export. This is reflected in the economics by 

an operating cost credit equal to this amount of electricity.

Figure 15. Distribution of plant electricity utilization by process area

3.9.3 Cost Estimation

The cooling tower was based on a cost estimate from Marley SPX for a fiberglass cooling tower 

capable of handling 44,000 gpm. Harris Group estimated costs for the cooling water circulation 

and makeup pumps using their internal database. The material of construction for the cooling 

water loop is carbon steel. The cost for the chiller came from a recent Trane quote for a 

similarly-sized system. Harris Group also used their internal database to estimate costs for the 

remaining equipment: the process water tank and pump; the plant/instrument air compressor, 

dryer, and surge tank; and the CIP system. 

to A100 

Feedstock 

Handling

2%

to A200 

Pretreatment

14%

to A300 

Hydrolysis & 

Fermentation

6%

to A400 

Enzyme 

Production

13% to A500 

Recovery

5%

to A600 

Wastewater

18%

to A700 Storage

0%

to A800 

Boiler/Turbogen

3%

to A900 Utilities

8%

Excess electricity 

to grid

31%

DW1102A

Total Power Generated = 41 MW
57 

4 Process Economics 

The ultimate purpose for developing such a detailed process design, simulation model, and cost 

estimate is to determine the economics of ethanol production. This information is used either as 

an absolute cost to assess ethanol’s potential in the marketplace or as a relative cost that can be 

used to guide research by examining the change in production cost associated with a process 

modification or other core research activity. 

The total capital investment (TCI) is first computed from the total equipment cost. Next, variable 

and fixed operating costs are determined. With these costs, we use a discounted cash flow 

analysis to determine the minimum ethanol selling price (MESP) required to obtain a zero net 

present value (NPV) with a finite internal rate of return (IRR). This section describes the 

assumptions made in completing the discounted cash flow analysis. 

Our analysis does not take into account any policy factors such as subsidies, mandates, or carbon 

credits because these would be purely speculative. The purpose of this analysis is to demonstrate 

whether or not cellulosic ethanol can be cost-competitive on its own merits and, if it cannot, to 

give policymakers a sense of the magnitude of incentive required to make it so. 

4.1 About Cost-Year Indices 

The cost-year of 2007 was chosen for this analysis to align with DOE’s 2012 cellulosic ethanol 

performance cost goals [27]. It is expected that these performance goals will remain in 2007 

dollars (2007$) through 2017 to permit comparison of future feedstocks, conversion 

technologies, and cellulosic biofuels. However, the present equipment costs were obtained in 

2009$ or 2010$ and cost-years for chemicals range from 1999 to 2010. The following is a brief 

discussion of the methods used for determining MESP in another year’s dollar value. 

Each piece of equipment has a purchased cost that reflects a quote made in a certain year. Capital 

costs provided by Harris Group in 2009$ or 2010$ were adjusted using the Plant Cost Index from 

Chemical Engineering Magazine [69] to a common basis year of 2007. The final cost index for a 

given year is generally not made available until the spring of the following year. Therefore, for 

the small number of equipment items that were quoted in 2010$, we used the Plant Cost Index 

value from August of 2010. Similarly, for chemical costs we used the Industrial Inorganic 

Chemical Index from SRI Consulting [70]. Employee salaries were obtained in 2009$ and were 

scaled using the labor indices provided by the United States Department of Labor Bureau of 

Labor Statistics [71]. The general formula for year-dollar back-casting is: 

2007 Cost = (Base Cost) ൬

2007 Cost Index

Base Year Index൰

The cost indices for capital and chemicals are not duplicated here; however, the reader is 

probably aware of the significant run-up in the price of raw materials such as steel and petroleum 

in the second half of the 2000s, beginning in 2003–2004. This inflation had dramatic effects on 

the prices of equipment and chemicals used in our design. Following a market correction, the 

Plant Cost Index crested in 2008, with the 2009 index actually being lower than the 2007 index. 

The Industrial Inorganic Chemical Index also crested in 2008 but the correction in 2009 was 

much smaller. In the year 2007 in particular, chemical prices varied wildly month-to-month and 
58 

location-to-location, depending on the available supply. It should be noted, therefore, that the 

chemical costs used in this analysis are static back-projections based on these cost indices. While 

these costs are reported in 2007$, they do not necessarily reflect the actual market price of that 

chemical in 2007.

The electricity export price was taken to be the average wholesale price as determined by the 

North American Electric Reliability Corporation for a given year. The wholesale price applies to 

electricity that is credited back to the grid. These data are provided on the Energy Information 

Administration’s website [72]. For 2007, the wholesale electricity price was 5.71675 cents/kWh.

4.2 Total Capital Investment (TCI) 

Section 3 of this report describes the details of the conceptual process design and how the 

purchased cost of the equipment was determined. The next step is to determine the installed cost 

of that equipment. The installation cost can be determined by performing a detailed study of 

everything required to install the necessary equipment and make it operational (e.g., foundation, 

piping, wiring). This type of detail is not warranted at this level of analysis, and a factored 

approach in which multipliers are applied to the purchased equipment cost is considered 

satisfactory. Many engineering texts such as Garrett [73] or Peters and Timmerhaus [74] detail 

such factored methodologies for common chemical processes. Indeed, there are as many 

methodologies for computing the total installed cost of capital equipment as there are textbooks 

on plant economics. The matter of which method is “correct” is of course open for debate;

however, at this level of detail, each methodology is generally considered as good as the next. 

The overarching cost methodology we have followed comes from the first biochemical design 

report prepared for NREL by Chem Systems in 1994 [75]. They provided installation factors for 

individual pieces of equipment, taken from Walas [76]. While simpler methods suggest a single 

multiplier be applied to the total purchased equipment cost, those methods tend to cite 

multipliers applicable for very large projects comprising mostly carbon steel equipment. For the 

relatively small plant described here, with mixed carbon and stainless steel equipment, the 

individual installation factor method makes more sense. 

From the 1994 Chem Systems report, some of the installation factors were modified in the 1999 

and 2002 design reports based on the experience of the engineers involved. Because the costs for 

the present design were all new, we chose to revert all installation factors back to the values 

specified by Walas. We found, however, that the Walas factors were in turn taken from the 

methodology of Cran [77], who deliberately excluded instrumentation costs from his installation

factors so that these could be estimated separately. Although Cran details how to compute 

instrument costs, we decided to assume these would be 30% of the purchased equipment cost, 

consistent with the estimate in Peters and Timmerhaus [74], which we imposed by adding 0.3 to 

the Cran factors. 

The list of installation factors for the various pieces of equipment is shown in Table 25. See 

Appendix A for a complete listing of the equipment, along with its purchased and installed costs. 

In contrast to the 2002 design, which had an equipment cost for every single item, Harris Group 

obtained many package quotes, in which a given unit operation and all of its support equipment 

were quoted under one price. The installation factor for such packages can be relatively low 

because most of the engineering is already included in the price. Additionally, equipment 
59 

designed as a pre-fabricated skid generally has a lower construction cost. Also, components that 

are more highly machined and have higher-quality metallurgy tend to be more expensive per unit 

mass and therefore have a lower installation factor as a function of purchase price than a less 

sophisticated component does. A good example of all three of these considerations is the 

pretreatment reactor system. 

Table 25. Installation Factors 

Equipment Multipliera

Agitators, carbon steel 1.6 

Agitators, stainless steel 1.5 

Boiler 1.8 

Compressors, motor driven 1.6 

Cooling tower 1.5 

Distillation columns, stainless steel 2.4 

Heat exchangers, shell & tube, stainless steel 2.2 

Heat exchangers, plate & frame, stainless steel 1.8 

Heat exchangers, air-cooled 2.8 

Inline mixers 1.0 

Skidded equipment 1.8 

Solids handling equipment (incl. filters) 1.7 

Pressure vessels, carbon steel 3.1 

Pressure vessels, stainless steel 2.0 

Pretreatment reactor system 1.5 

Pumps, stainless steel 2.3 

Pumps, carbon steel 3.1 

Tanks, field-erected, carbon steel 1.7 

Tanks, field-erected, stainless steel 1.5 

Tanks, storage, plastic 3.0 

Tanks, storage, carbon steel 2.6 

Tanks, storage, stainless steel 1.8 

Turbogenerator 1.8 a Installed cost = (purchased equipment cost) x (multiplier). 

The purchased cost for a given component reflects a baseline equipment size. As changes are 

made to the process, the equipment size required may be different than what was originally 

designed. Instead of re-costing in detail, an exponential scaling expression was used: 

New Cost ൌ ሺBase Costሻ ൬

New Size

Base Size൰

௡

where n is a characteristic scaling exponent (typically in the range of 0.6 to 0.7) based upon 

some characteristic of the equipment related to production capacity, such as flow or heat duty. 

Such scaled costs are easier to calculate and generally give nearly the same result as resizing the 

equipment for each scenario. The scaling exponent can be inferred from vendor quotes if 

multiple quotes are given for different sizes, computed with costing software, or obtained from a 

standard reference [73–74, 76]. In our case, we returned to the exponents proposed in the 1994 

Chem Systems report, which are shown in Table 26. Note that because the costs for most 

equipment items were newly estimated at their size in the present design, the exponents do not 

greatly affect the base MESP. They only factor into the calculation when changes are made in 

the process. 
60 

Table 26. Scaling Exponents 

Equipment Exponent

Agitators 0.5

Compressors, motor driven 0.6

Distillation columns 0.6

Heat exchangers 0.7

Inline mixers 0.5

Package quotes / Skidded equipment 0.6

Pressure vessels 0.7

Pumps 0.8

Tanks, atmospheric 0.7

Solids handling equipment 0.8

Once the total equipment cost has been determined in the year of interest, we must add several 

other direct and indirect costs to determine the total capital investment (TCI). Site development 

and warehouse costs are based on the inside-battery-limits (ISBL) equipment costs (Areas 200, 

300, 400, and 500) and are considered part of the total direct cost (TDC). Project contingency, 

field expenses, home-office engineering and construction activities, and other costs related to 

construction are computed relative to the TDC and give the fixed capital investment (FCI) when 

summed. The sum of FCI and the working capital for the project is the TCI. Table 27

summarizes these categories and additional factors. Harris Group reviewed the items and factors 

and determined that they were consistent with other projects the company has completed. Harris 

Group thought the “home office and construction” charge used in 2002 was possibly a bit high at 

25%, so this was reduced to 20%. The project contingency was also changed from 3% to 10% of 

the TDC. Table 28 summarizes the total installed equipment cost and illustrates the application 

of factors in Table 27 to obtain the total capital investment (TCI). 
61 

Table 27. Additional Costs for Determining Total Capital Investment (TCI)

Item Description Amount

Additional direct costs

Warehouse On-site storage of equipment and supplies. 4.0% of the 

installed cost of 

ISBL equipment 

(A200,300,400,500) 

Site development Includes fencing, curbing, parking lot, roads, well 

drainage, rail system, soil borings, and general paving. 

This factor allows for minimum site development 

assuming a clear site with no unusual problems such as 

right-of-way, difficult land clearing, or unusual 

environmental problems. 

9% of ISBL

Additional piping To connect ISBL equipment to storage and utilities 

outside the battery limits.

4.5% of ISBL

Indirect costs

Prorateable costs This includes fringe benefits, burdens, and insurance of 

the construction contractor. 

10% of total direct

cost (TDC)

Field expenses Consumables, small tool and equipment rental, field 

services, temporary construction facilities, and field 

construction supervision.

10% of TDC

Home office and 

construction 

Engineering plus incidentals, purchasing, and

construction. 

20% of TDC

Project contingency Extra cash on hand for unforeseen issues during 

construction.

10% of TDC

Other costs Start-up and commissioning costs. Land, rights-of-way, 

permits, surveys, and fees. Piling, soil 

compaction/dewatering, unusual foundations. Sales, use, 

and other taxes. Freight, insurance in transit, and import 

duties on equipment, piping, steel, instrumentation, etc. 

Overtime pay during construction. Field insurance. Project 

team. Transportation equipment, bulk shipping containers, 

plant vehicles, etc. 

10% of TDC
62 

Table 28. Project Cost Worksheet Including Total Direct Costs and Total Capital Investment

Process Area Purchased Cost Installed Cost

 Area 100: Feedstock handlinga $ 14,200,000 $ 24,200,000

 Area 200: Pretreatment $ 19,900,000 $ 29,900,000

 Area 200: Conditioning $ 1,500,000 $ 3,000,000

 Area 300: Enzymatic hydrolysis and fermentation $ 18,500,000 $ 31,200,000

 Area 400: Enzyme production $ 10,700,000 $ 18,300,000

 Area 500: Recovery $ 11,100,000 $ 22,300,000

 Area 600: Wastewater $ 49,300,000 $ 49,400,000

 Area 700: Storage $ 2,800,000 $ 5,000,000

 Area 800: Boiler $ 36,500,000 $ 66,000,000

 Area 900: Utilities $ 4,000,000 $ 6,900,000

Totals (excl. Area 100) $ 154,500,000 $ 232,100,000

 Warehouse 4.0% of ISBL $ 4,200,000

 Site development 9.0% of ISBL $ 9,400,000

 Additional piping 4.5% of ISBL $ 4,700,000

Total Direct Costs (TDC) $ 250,400,000

 Prorateable expenses 10.0% of TDC $ 25,000,000

 Field expenses 10.0% of TDC $ 25,000,000

 Home office & construction fee 20.0% of TDC $ 50,100,000

 Project contingency 10.0% of TDC $ 25,000,000

 Other costs (start-up, permits, etc.) 10.0% of TDC $ 25,000,000

Total Indirect Costs $ 150,200,000

Fixed Capital Investment (FCI) $ 400,600,000

 Land $ 1,800,000

 Working capital 5.0% of FCI $ 20,000,000

Total Capital Investment (TCI) $ 422,500,000

Lang Factor (FCI/purchased equip cost)b 3.1

TCI per annual gallon $6.92/gal

2007 dollars

DW1102A a Feedstock handling not included in this calculation. b Area 600 not included in Lang Factor. 

4.3 Variable Operating Costs 

Variable operating costs, which include raw materials, waste handling charges, and by-product 

credits, are incurred only when the process is operating. Quantities of raw materials used and 

wastes produced were determined using the Aspen material balance. Table 29 documents the 

costs and sources of chemicals used in the process, and Table 30 summarizes the variable costs

on a per-year and per-gallon-of-ethanol basis.
63 

Table 29. Chemical Costs and Sources 

Component Cost (2007$) Source

Feedstock (corn stover) $0.0224/lb MYPP, $58.50/dry ton @ 20% moisture [27]

Sulfuric acid, 93% $0.0399/lb Basic Chemical of Omaha via Harris Group

Ammonia $0.1993/lb Terra Industries via Harris Group

Corn steep liquor $0.0252/lb Corn Products via Harris Group

Diammonium phosphate $0.4385/lb Ronas Chemicals via Harris Group

Sorbitol $0.5005/lb Coast Southwest via Harris Group

Glucose $0.2579/lb USDA ERS [78]

SO2 $0.1378/lb SRI Chemical Economics Handbook (CEH)

Enzyme nutrients $0.3727/lb SRI CEH (See Table 20)

Caustic $0.0678/lb Brown and Caldwell WWT design [65]

Lime $0.0904/lb Harris Group

Boiler chemicals $2.2661/lb 2002 Design Report [2]

Cooling tower chemicals $1.3580/lb 2002 Design Report [2]

Fresh water $0.0001/lb Peters & Timmerhaus [74]

Table 30. Variable Operating Costs

Process 

Area

Stream Description Usage

(kg/h)

Usage 

(lb/h)

Cost 

($/ton)

$/hour MM$/yr 

(2007$)

Cent/Gal 

Ethanol 

(2007$)

Raw materials

N/A Feedstock (wet) 104,167 229,688 46.80 5,374.69 45.20 74.07

A200 Sulfuric acid, 93% 1,981 4,367 81.39 177.73 1.49 2.45

Ammonia 1,051 2,317 406.96 471.48 3.97 6.50

A300 Corn steep liquor 1,158 2,554 51.55 65.84 0.55 0.91

Diammonium phosphate 142 313 895.32 140.33 1.18 1.94

Sorbitol 44 98 1,021.93 49.96 0.42 0.69

A400 Glucose 2,418 5,332 526.52 1,403.60 11.80 19.34

Corn steep liquor 164 363 51.55 9.35 0.08 0.13

Ammonia 115 254 406.96 51.59 0.43 0.71

Host nutrients 67 149 745.30 55.34 0.47 0.76

Sulfur dioxide 16 36 275.70 4.99 0.04 0.07

A600 Caustic 2,252 4,966 135.65 336.83 2.83 4.64

A800 Boiler chems <1 1 4,532.17 1.23 0.01 0.02

FGD lime 895 1,973 180.87 178.42 1.50 2.46

A900 Cooling tower chems 2 5 2,716.10 7.14 0.06 0.10

Makeup water 147,140 324,443 0.23 38.11 0.32 0.53

Subtotal 8,328.49 70.36 115.30

Waste disposal

A800 Disposal of ash 5,725 12,623 28.86 182.15 1.53 2.51

Subtotal 182.15 1.53 2.51

By-products and credits

Grid electricity 12,797 kW $0.0572/kWh 731.57 6.15 10.08

Area 100 electricity 859 kW $0.0572/kWh 49.13 0.41 0.68

Subtotal 780.71 6.57 10.76

Total variable operating costs 7,779.08 65.33 107.76
64 

4.4 Fixed Operating Costs

Fixed operating costs are generally incurred in full whether or not the plant is producing at full 

capacity. These costs include labor and various overhead items. Many of the assumptions on 

fixed operating costs follow the 2002 design report [2] and/or Peters and Timmerhaus [74]. 

Table 31 shows the recommended number of employees and associated salaries. The number of 

employees was estimated by considering the likely degree of automation for each area and 

adding a reasonable number of management and support employees. The total number of 

employees has been reduced from 77 in the 2002 design to 60. The major reason for the reduced 

labor requirement is the redesign of the feed handling area. In the present design, feedstock is 

received year-round according to uniform-format specifications and is unloaded with truck 

tippers, eliminating the need for yard employees and operators to perform milling and washing. 

Eight operators and two laboratory technicians were added to support the on-site enzyme section, 

but overall fewer employees were required. Excluding these 10 employees, 50 employees for the 

ethanol facility is also consistent with a 2008 LECG consulting report prepared for the 

Renewable Fuels Association [79]. Salaries were estimated by using commercially available 

salary estimators such as Salary.com as well as open positions posted on internet job sites. 

Because the model feedstock is corn stover, salaries were estimated for rural regions of the U.S.

Midwest (Iowa, Missouri, etc.). These estimates may vary depending on location. 

Table 31. Fixed Operating Costs

Position 2007 

Salary

# 

Required

2007 Cost MM$/yr

(2007$)

Cent/Gal 

Ethanol 

(2007$)

Labor and supervision

Plant manager 141,569 1 141,569

Plant engineer 67,414 2 134,828

Maintenance supervisor 54,894 1 54,894

Maintenance technician 38,522 12 462,266

Lab manager 53,931 1 53,931

Lab technician 38,522 2 77,044

Lab tech-enzyme 38,771 2 77,044

Shift supervisor 46,227 4 184,906

Shift operators 38,522 20 770,443

Shift oper-enzyme 38,771 8 308,177

Yard employees 26,966 4 107,862

Clerks and secretaries 34,670 3 104,010

Total salaries 2,500,000 2.48 4.06

Labor burden (90%) 2,200,000 2.23 3.65

Other overhead

Maintenance 3.0% of ISBL 3,200,000 3.14 5.15

Property insurance 0.7% of FCI 2,700,000 2.80 4.60

Total fixed operating costs 10.66 17.46

A 90% labor burden is applied to the salary total and covers items such as safety, general 

engineering, general plant maintenance, payroll overhead (including benefits), plant security, 

janitorial and similar services, phone, light, heat, and plant communications. The 90% estimate is 

the median of the general overhead range suggested in the 2008 PEP Yearbook produced by SRI 

Consulting [80]. 
65 

Annual maintenance materials were estimated as 3% of the installed ISBL capital cost and 

property insurance and local property tax were estimated as 0.7% of the fixed capital investment, 

based on the 1994 Chem Systems report [75].

4.5 Discounted Cash Flow Analysis and the Minimum Selling Price of Ethanol

Once the total capital investment, variable operating costs, and fixed operating costs have been 

determined, a discounted cash flow rate of return (DCFROR) analysis can be used to determine 

the minimum selling price per gallon of ethanol produced. The discounted cash flow analysis is 

calculated by iterating the selling cost of ethanol until the net present value of the project is zero. 

This analysis requires that the discount rate, depreciation method, income tax rates, plant life, 

and construction start-up duration be specified. Because this plant is equity-financed, some 

assumptions about the loan terms are also required.

We stress that the financial assumptions detailed in this section are of course subject to 

uncertainty and are not intended to represent the final word in project management. For research 

and policy purposes, one must at some point decouple the cost of making ethanol from the cost 

of making money. These financial parameters serve merely as a reference point from which to 

examine other economic sensitivities.

Discount Rate

For this analysis, the discount rate (which is also the internal rate of return [IRR] in this analysis) 

was set to 10% and the plant lifetime was set to 30 years. The discount rate was also used in 

previous design reports and was based on the recommendation in Short et al. [81] on how to 

perform economic evaluations of renewable energy technologies for DOE. His view was that, 

“In the absence of statistical data on discount rates used by industrial, transportation and 

commercial investors for investments with risks similar to those of conservation and renewable 

energy investments, it is recommended that an after tax discount rate of 10%...be used.” 

Equity Financing

For this analysis, it was assumed that the plant would be 40% equity financed. The terms of the 

loan were taken to be 8% interest for 10 years. The principal is taken out in stages over the 3-

year construction period. Interest on the loan is paid during this period, but principal is not paid 

back. (This is another n

th-plant assumption, which says that this cash flow comes from the parent 

company until the plant starts up.) This is all in contrast to the 2002 design, which assumed a 

100% equity purchase. Figure 16 illustrates the sensitivity of minimum ethanol selling price to 

the percentage of equity financing and the after-tax discount rate (the IRR). 
66 

Figure 16. Sensitivity of MESP to IRR and % equity (8% interest on a 10-year loan)

Depreciation

To determine the capital depreciation amount for the calculation of federal taxes to be paid we 

used the IRS Modified Accelerated Cost Recovery System (MACRS). Within the MACRS 

system is the General Depreciation System (GDS), which allows both the 200% and 150% 

declining balance (DB) methods of depreciation. This offers the shortest recovery period and the 

largest tax deductions. According to IRS publication 946 [82], a cellulosic ethanol plant would 

fall under Asset Class 49.5, “Waste Reduction and Resource Recovery Plants.” This class uses a 

7-year recovery period, not including the steam plant equipment, which has a 20-year recovery 

period (Asset Class 49.13). IRS publication 946 contains a special provision for cellulosic 

biofuels plants that allows them to write off 50% of the capital investment in the first year. This 

was not implemented in our cost model because it does not ultimately affect the MESP. 

(Although the provision affects the cash flow in the first few years of the analysis, it does not 

change the year in which the plant goes into the black and must start paying taxes.)

Taxes

The federal corporate tax rate used in our analysis is 35%. Income tax is averaged over the plant 

life and that average is calculated on a per-gallon basis. The amount of income tax to be paid by 

a potential ethanol producer varies annually due to changes in the volume of product produced 

and the allowable depreciation deduction. In fact, no income tax is paid in the first eight years of 

operation because the depreciation and loan interest deductions are greater than the net income. 

State taxes are not considered, primarily because the location of the plant has not been 

determined and tax rates vary from state to state (from 0% to 12%).

Construction Time

The construction time is important to the cash flow analysis because no income is earned during 

construction, but huge sums of money are being expended. Perry and Green [83] indicate that 

$2.00

$2.20

$2.40

$2.60

$2.80

$3.00

$3.20

$3.40

$3.60

0% 20% 40% 60% 80% 100%

Minimum Ethanol Selling Price ($/gal)

Equity %

10% IRR

15% IRR

20% IRR

DW1102A
67 

small projects (less than $10 million investment) can be constructed in fewer than 18 months and 

that larger projects can take up to 42 months. An overview of petroleum refining economics 

indicates that large refineries (on the order of $1.5 billion investment) can be constructed in 24 

months [84]. Certainly this ethanol process is much smaller than a petroleum refinery, so using a 

construction time of 24 months fits within these references, although an important difference 

between this type of facility and a refinery is the large number of field-erected vessels. These are 

constructed on-site and have a longer construction time than if the tanks were delivered finished. 

Table 32 summarizes the schedule for construction and the cash flow during that time. Twelve

months are added before construction for planning and engineering. 

Table 32. Construction Activities and Cash Flow

Project 

Start 

Month

Project 

End 

Month

Activity Description 

% of 

Project 

Cost

0 12 Project plan and schedule established; conceptual and basic 

design engineering, permitting completed. Major equipment bid 

packages issued, engineering started on selected sub-packages, 

P&IDs complete, preliminary plant and equipment arrangements 

complete. 

8%

12 24 All detailed engineering including foundations, structure, piping, 

electrical, site, etc. complete; all equipment and instrument 

components purchased and delivered; all site grading, drainage, 

sewers, rail, fire pond, foundation, and major structural 

installation complete; 80% of all major process equipment set (all 

except longest-lead items), all field fabricated tanks built, and the 

majority of piping and electrical materials procured. 

60%

24 36 Complete process equipment setting, piping, and instrumentation 

installation complete; all electrical wiring complete; all building 

finishing and plumbing complete; all landscaping complete; precommissioning complete; and commissioning, start-up, and initial 

performance test complete. 

32%

TOTAL 100%

Note: The above assumes no utility or process equipment orders placed prior to month seven. Expenditures based on 

typical 60 MMgal/yr grain-to-ethanol facility.

Start-Up Time

Perry and Green [83] indicate that for a moderately complex plant, start-up should be about 25% 

of the construction time, or 6 months in this case. Delta-T’s experience (described in the 2002 

design report) with start-up indicated that a large grain-to-ethanol plant could be started up in 

less than 6 months. Considering that this design is for the n

th operating plant, we assumed a startup time of 3 months. The start-up period is not completely wasted, however. We expect that an 

average of 50% production could be achieved during that period while incurring 75% of variable 

expenses and 100% of fixed expenses. 

Working Capital

Peters and Timmerhaus [74] define working capital as money available to cover (1) raw 

materials and supplies in inventory, (2) finished product in storage, (3) accounts receivable, (4) 

cash on hand for monthly payments such as wages and maintenance supplies, (5) accounts 

payable, and (6) taxes payable. They indicate that working capital is usually 10%–20% of the 

fixed capital investment. This flow of money is required over the life of the plant, beginning in 
68 

the start-up phase to make product that generates revenue to use in purchasing more materials 

and supplies. For this project, 10% would be about $40 million. On-site ethanol storage capacity 

is 7 days; assuming that the product can be made, shipped, and payment received in 30 days is 

conservative. One month’s raw materials, labor, maintenance, taxes, and overhead is $6.9 

million. Therefore, a lower number seems reasonable. Garrett [73] suggests that using a fraction 

of the yearly operating cost, typically 10%–35%, is more relevant. When these percentages are 

used to calculate working capital, this results in a working capital range of $8 million–$28 

million. We chose to use $20 million, or 5% of the fixed capital investment.

Table 33 summarizes the parameters used in the discounted cash flow analysis. Using these 

parameters, plus the cost information in Table 28, Table 30, and Table 31, the resulting 

minimum ethanol selling price of pure ethanol is $2.15/gal (2007$). Table 34 summarizes the 

yields and conversion costs for the present design. According to the methodology of Cran [77], 

the expected accuracy of the TCI analysis is +/- 25%. If we apply this uncertainty to the TCI, the 

impact on the cost of ethanol is +/-$0.24/gal. The complete discounted cash flow summary 

worksheet is shown in Appendix B. The MESP can be further broken down into the cost of each 

process area. Figure 17 illustrates the contribution to the overall cost by process area and capital, 

operations, and fixed costs. (The bar for feedstock + handling reflects the single feedstock cost of 

$58.50/dry U.S. ton delivered to pretreatment and has not been decomposed.) 

Table 33. Discounted Cash Flow Analysis Parameters

Plant life 30 years

Discount rate 10%

General plant depreciation 200% declining balance (DB)

General plant recovery period 7 years

Steam plant depreciation 150% DB

Steam plant recovery period 20 years

Federal tax rate 35%

Financing 40% equity

Loan terms 10-year loan at 8% APR

Construction period 3 years

 First 12 months’ expenditures 8%

 Next 12 months’ expenditures 60%

 Last 12 months’ expenditures 32%

Working capital 5% of fixed capital investment

Start-up time 3 months

 Revenues during start-up 50%

 Variable costs incurred during start-up 75%

 Fixed costs incurred during start-up 100%
69 

Table 34. Summary of Yields, Rates, and Conversion Costs

Feedstock rate 2,205 dry U.S. ton/day

On-line time 8,410 h/yr

Ethanol yield 79.0 gal/dry U.S. ton feedstock

(76% theoretical based on feedstock carbohydrates)

Ethanol production rate 61.0 MMgal/yr

Total equipment cost $232MM

Total capital investment (TCI) $423MM

TCI per annual gallon $6.92/gal

Minimum Ethanol Selling Price $2.15/gal

 Feedstock contribution $0.74/gal

 Enzyme contribution $0.34/gal

 Non-enzyme conversion contribution $1.07/gal

Figure 17. Cost contribution details from each process area (per gallon EtOH)

At the time of the 2002 design report, it was thought that the costs predicted then were probably 

too high for a company to start constructing a cellulosic ethanol plant based on undemonstrated 

technology. Today we are seeing several such biorefineries being planned or built at the pilot and 

commercial scale. The technoeconomic results reported here fall within the range claimed by 

these companies with respect to ethanol yield and TCI/gallon. The principal barrier to 

commercialization today is the current lack of access to capital for large, first-of-a-kind projects. 

-$0.40 -$0.20 $0.00 $0.20 $0.40 $0.60 $0.80

Feedstock + Handling: 74.1¢ 

Pretreatment / Conditioning: 28.8¢ 

Enzymatic Hydrolysis and Fermentation: 20.2¢ 

Cellulase Enzyme: 33.8¢ 

Distillation and Solids Recovery: 12.1¢ 

Wastewater Treatment: 33.5¢ 

Storage: 2.3¢ 

Boiler/Turbogenerator: (Net) 4.3¢ 

Utilities: 6.3¢ 

Capital Recovery Charge Raw Materials & Waste Process Electricity

Grid Electricity Total Plant Electricity Fixed Costs

DW1102A
70 

5 Analysis and Discussion 

5.1 Carbon and Energy Balance

Overall material and energy balances were performed on the Aspen model. Table 35 shows the 

flow of carbon inputs and outputs. Closure on the carbon balance was 100.4%; small errors are 

due to the imperfect stoichiometry used to generate microbial cell mass in fermentation and 

wastewater treatment. As shown in the table, >96% of all carbon in the process enters in the 

biomass feed, with small amounts of additional carbon coming from glucose (for enzyme 

production) and fermentation nutrients like corn steep liquor. 

Of the carbon inputs to the process, 29% of carbon atoms leave as ethanol. Major exit points for 

the balance of carbon (accounting for 66%) are the combustor stack and scrubber vent. We 

expect the amount of carbon exiting in the combustion exhaust to be rather large because most 

by-products (lignin, etc.) of this process are burned to form CO2. The majority of the carbon 

exiting the scrubber vent is also in the form of CO2 from fermentation. 

Table 35. Ethanol Plant Overall Carbon Balance

Stream Carbon Flow 

(kmol/h)

% of Carbon 

Flow

Carbon inlets

Stover feedstock 3,117 96.6%

Glucose 81 2.5%

Other 28 0.9%

Total 3,226 100.0%

Carbon outlets

Combustion exhaust 1,670 51.8%

Ethanol product 941 29.2%

Scrubber vent 471 14.6%

Ash 10 0.3%

WWT brine 4 0.1%

Aerobic lagoons 88 2.7%

Enzyme production vents 54 1.7%

Total 3,238 100.4%

Only the carbohydrate components of the biomass can be converted to ethanol by fermentation. 

Of the carbon in the feed stream, 1,857 kmol/h (59.6%) is carbohydrates (glucan, xylan, 

arabinan, mannan, galactan, and sucrose); converting all of this carbon to ethanol represents the 

theoretical yield. According to Table 35, 1,412 kmol/h of the feed stream carbon is actually 

converted by the ethanol pathway (ethanol in the product plus CO2 in the scrubber vent), so the 

actual yield is 76% of the theoretical yield. The same yield in the 2002 design was 80% of 

theoretical. Ours is slightly lower because we have not assumed conversion of the minor sugars 

galactose and mannose.

An energy balance was performed to compare the energy products from the process (ethanol and 

electricity) to the potential energy in the biomass feed. The potential energy for material streams 

used in this analysis has three contributions: 

1. A combustion energy, defined as the lower heating value (LHV) of the stream at a 

reference temperature of 25°C. 
71 

2. A sensible heat effect to account for deviations of a stream’s temperature from the 

reference temperature of 25°C. (This is computed by Aspen.) 

3. A latent heat effect to account for major volatile components (e.g., water, ethanol, 

furfural) that are in a phase other than the reference phase. 

The LHV is the energy released when a material is burned to form its combustion products, with 

water in the vapor phase. This is considered to be a more “true” heating value for transportation 

fuels, where the combustion water is generally not condensed to recover heat. On an LHV basis,

any liquid water present in a material stream therefore reduces its combustion potential because 

heat must be expended to vaporize that water. Note that the sign convention in this basis follows 

the usual convention for heat of combustion: streams that can be burned (e.g., glucose or ethanol) 

have a negative potential energy. Streams that cannot be burned (e.g., water or sulfuric acid) 

have a positive potential energy. 

A comparison of the major energy streams produced by the process to the combustion energy of 

the biomass feed is made in Table 36; a small discrepancy in the model’s energy balance closure 

is due to reasons cited above. The combined energy available from the ethanol product and the 

exported electricity is 48% of the energy content of the biomass feed. The energy losses from the 

cooling tower and combustion exhaust are fairly large. Ambient heat losses are also significant; 

these include the wastewater lagoons and the air-cooled condenser on the rectification column. 

Compared to the 2002 design, less energy is lost from the cooling tower and combustion exhaust 

but more is lost through these ambient mechanisms. Note that “electricity reserve” refers to the 

total power demand to a small number of intermittent users in the plant (e.g., conveyor belts). 

This reserved electricity is not available for export and must be accounted for separately. 

Table 36. Ethanol Plant Overall Energy Balance

Stream Energy 

Flow 

(MMkcal/h)

Major energy inputs

Stover feedstock -316

Glucose -8

Sulfuric acid 1

Ammonia -4

Other chemicals 0

Well water 87

Total -239

Major energy outputs

Ethanol -139

Cooling tower evaporation -20

Combustion exhaust + ash -12

Ambient heat losses -63

Electricity export -11

Electricity reserve -4

WWT brine 8

Total -240

As with the carbon balance, the energy content of non-carbohydrate compounds cannot be 

converted to ethanol. The total energy content of the carbohydrate components in the feed is -192 
72 

MMkcal/h. In other words, only 61% of the feedstock’s combustion energy can be used to make 

ethanol. Given the energy content recovered as ethanol from Table 36, the percent of 

carbohydrate energy in the feed recovered as ethanol is about 72%, similar to the amount of 

carbohydrate carbon recovered as ethanol. A diagram of the overall mass, carbon, and energy 

balance is presented in Figure 18.
73 

Figure 18. Mass, carbon, and energy balance diagram for the overall process

Pretreatment

Chemicals

Biomass

Enzymatic 

hydrolysis

Chemicals

Glucose

Enzyme 

Production

Wastewater 

treatment

Boiler/

Turbogenerator

-147 Recovery

Process and

Cooling water

Stack+Ash

Ethanol

Well water

Cooling twr

WWT Evap

Elec out

losses

Fermentation

Chemicals

Chemicals

Vent

Air

Chemicals

Air

Air

Air

Elec Rsrv

Brine

losses

losses

1,299 kg/hr

22 kmol C/hr

-2 Gcal/hr

3,031 kg/hr

0 kmol C/hr

-3 Gcal/hr

104,167 kg/hr

3,117 kmol C/hr

-316 Gcal/hr

12,105 kg/hr

2,845 kg/hr

81 kmol C/hr

-8 Gcal/hr

377 kg/hr

6 kmol C/hr

-1 Gcal/hr

32,853 kg/hr

223,602 kg/hr

4,504 kg/hr

0 kmol C/hr

2 Gcal/hr

895 kg/hr

0 kmol C/hr

0 Gcal/hr

284,495 kg/hr

147,140 kg/hr

0 kmol C/hr

87 Gcal/hr

817,045 kg/hr

3,226 kmol C/hr

-239 Gcal/hr

21,398 kg/hr

471 kmol C/hr

0 Gcal/hr

Vent

21,808 kg/hr

941 kmol C/hr

-139 Gcal/hr

-24 Gcal/hr

33,387 kg/hr

54 kmol C/hr

0 Gcal/hr

221,417 kg/hr

88 kmol C/hr

-1 Gcal/hr

9,929 kg/hr

4 kmol C/hr

8 Gcal/hr

-26 Gcal/hr

-11 Gcal/hr

-12 Gcal/hr

-4 Gcal/hr

371,750 kg/hr

1,680 kmol C/hr

-12 Gcal/hr

137,362 kg/hr

0 kmol C/hr

-20 Gcal/hr

817,052 kg/hr

3,238 kmol C/hr

-240 Gcal/hr

-111

17

4

80

86

7

-97

-66

216

-51

-3

-4 -39

-1

-3

-1

-6

101

Lignin cake

Thin stillage

Recycle water

Biogas

Recycle water

Dilution water

Dilution waterDilution water

Cooling water

Material Heat Work Energy flows (Gcal/hr) shown on selected intermediate streams.

Positive energy flow has a negative sign.

Evaporation
74 

5.2 Water Balance

Consumptive water usage in the biochemical conversion process is generally considered to be 

high compared with a thermochemical conversion process or even a petroleum-based fuel

process, when viewed in terms of gallons of water used per gallon of fuel produced (gal/gal). 

Based on detailed water balances from the Aspen simulation, the consumptive water usage in the 

2002 design was 6.3 gallons of water used per gallon of ethanol [2]. Using the older model as a 

guide, Harris Group adjusted some assumptions around the cooling tower for a modest reduction 

in water usage, but as the model was updated for the present report, these adjustments were 

ultimately canceled out by more accurate accounting of cooling water streams and inclusion of 

the on-site enzyme production area. Without additional process modifications, the new design’s 

consumptive water usage would therefore have been greater than in the 2002 design. In light of 

this, we gave more attention to water reduction and optimization. 

As discussed in Section 3, the switch from lime to ammonia conditioning after pretreatment 

created a problem with ammonium sulfate and other inorganics accumulating in the evaporated 

stillage syrup. Boiler vendors were not comfortable quoting equipment to handle this stream, so 

we chose instead to handle the stillage in the wastewater treatment section. Doing so gave us the 

opportunity to remove the syrup evaporators, which in the 2002 design accounted for about onehalf of the cooling tower losses and one-third of the total water losses. (Although the evaporators 

returned clean water to the process in the 2002 design, they were a net water loser—the amount 

of water returned to the process was less than the amount of water lost from the cooling tower as 

a result of condensing the recycled water.) Furthermore, the evaporated stillage syrup, which was 

sent to the combustor in the 2002 design, contained about 50% water. This water came out of the 

ethanol conversion process but was then lost via the combustor stack. By routing the stillage to 

wastewater treatment, all of the water contained in the stillage becomes theoretically recyclable 

to the process. The stillage evaporators in the 2002 design also acted as the condenser for the 

rectification column. Simply removing the evaporators and using cooling water in the rectifier 

condenser would therefore not have resulted in a significant reduction in consumptive water 

usage. Instead, Harris Group proposed placing an air-cooled condenser on the rectification 

column. The capital cost was about the same—$487K for the air-cooled condenser versus $440K 

for the cooling-water condenser initially specified by Megtec. The air-cooled unit uses about 300 

hp (225 kW) of electricity, but this extra demand is insignificant to the overall plant, which 

generates more than 40 MW of electricity.

Taken together, elimination of the stillage evaporators combined with addition of the air-cooled 

rectifier condenser and improved heat integration in the front end of the process led to a 

reduction in the consumptive water usage from 6.3 gal/gal in 2002 to 5.4 gal/gal in the present 

design report. The overall water balance is shown in Table 37—the “makeup water” line item 

indicates the consumptive water usage. 

Additional water enters the process via the feedstock and raw chemicals, both as free water and 

as “potential” water, i.e., the combustion product of lignin and unconverted sugars. Water is also 

consumed in hydrolysis reactions, and these are accounted for as well. We note that describing 

water usage with the units of gallons of water per gallon of ethanol can be somewhat misleading, 

because this quantity obviously depends on ethanol yield, and the yield in the present design is 

about 12% lower than in the 2002 design (79.0 gal/ton versus 89.7 gal/ton). Based on the actual

makeup water flow, the present design has reduced the 2002 water usage by about 20% (150,000 
75 

kg/h versus 187,000 kg/h) for a plant that has the same feedstock flow rate and produces its own 

enzyme.

Table 37. Ethanol Plant Water Balance

Inputs kg/h gal/gal Outputs kg/h gal/gal

Moisture in feedstock 20,833 0.76 Water in ethanol product 109 0.00

Water in glucose (Area 400) 427 0.02 Cooling tower evaporation 137,363 5.00

Water in raw chemicals 3,056 0.03 Stripped in enzyme aeration 774 0.03

Generated in enzyme prod. 1,291 0.02 Scrubber vent 522 0.02

Generated in fermentation 121 0.00 Destroyed in pretreatment 2,473 0.09

Generated in WWT 2,680 0.10 Destroyed in enzym. hydrolysis 2,631 0.10

Generated in combustor 23,869 0.87 WWT evaporation 4,350 0.16

Enzyme prod. air intake 638 0.02 WWT brine 4,967 0.18

Lignin cake dryer intake 237 0.01 Combustor stack 54,451 1.98

WWT lagoon aeration intake 4,379 0.16 Boiler blowdown vent 2,580 0.09

Combustor air intake 5,572 0.20

Makeup (consumptive) water 147,140 5.35

Sum of inputs 210,243 7.65 Sum of outputs 210,220 7.65

The ethanol conversion process itself is not a significant loser of water. Utilities are responsible 

for practically all water loss, with 90% of the net water loss due to cooling tower evaporation. 

Table 38 expresses the individual cooling water users (see Figure 14) in terms of their water loss 

responsibility in gal/gal. 

Table 38. Individual Contributors to Cooling Water Evaporation 

Cooling Water User MMkcal/h % of 

duty

gal/gal

Condensing turbine 50.7 52.1% 2.60

Chiller condenser 13.4 13.8% 0.69

Pre-hydrolysis cooler 10.6 10.9% 0.54

Hydrolysate flash 7.3 7.5% 0.37

Pre-fermentation cooler 5.9 6.1% 0.30

Ethanol product cooler 5.2 5.3% 0.27

Molsieve regeneration cooler 2.6 2.6% 0.13

Enzyme air compressor 1.2 1.2% 0.06

Beer column condenser 0.5 0.5% 0.02

Total 97.4 100% 5.00

The largest cooling water user is the condensing turbine (M-811), which is responsible for 2.6 

gal/gal of water loss. A condenser on the turbine allows the steam to discharge from M-811 at a 

vacuum (about 0.1 atm) for maximum conversion of compressive energy to electricity. With the 

condensing turbine in place, the process generates enough electricity to supply all users in the 

plant plus a significant surplus, which is assumed to be sold to the grid. It would be difficult to 

eliminate the condensing turbine in the present design because without it, the plant would only 

be able to burn enough of the lignin and biogas to meet the steam demand. This would result in a 

significant excess of fuel that would have to be landfilled or disposed of somehow, or incinerated

at no value. The plant would also no longer make enough electricity to support all of its users 

and would have to purchase the balance from the grid. Note that grid electricity also has a water 

usage associated with it, which in a full life cycle assessment would have to be attributed to the 

ethanol plant. Given the turbogenerator electricity output, the consumptive water use associated 
76 

with the condensing turbine can be expressed as 0.46 gal/kWh, which is roughly equivalent to 

the consumptive water use in coal-fired electricity plants [85].

The water-condensed chiller assumed here could potentially be replaced with an air-condensed 

system; however, these units do not have a very large capacity, so many parallel units would be 

required at a significant capital expense. The hydrolysate flash condenser and the pre-hydrolysis 

cooler could also potentially be replaced, but by inspection neither of them would break even 

with their water savings given the low cost of water used in the model. 

The three low-temperature coolers in Table 38 (ethanol product cooler, pre-fermentation cooler, 

and molecular sieve regeneration cooler) account for another ~0.7 gal/gal of water loss, but these 

are not candidates for air-cooled heat exchangers because their outlet temperature is too low 

(32°–35°C). The other users below these in Table 38 either do not represent a significant cooling 

water duty or are likely to come skidded with cooling water hookups. 

5.3 The Cost of Sugar

Sugars are an intermediate product in the formation of ethanol. Hemicellulose is hydrolyzed to 

form xylose, arabinose, galactose, and mannose in the pretreatment area. Cellulase enzyme 

hydrolyzes the cellulose into glucose. NREL has received feedback from the biofuels industry 

(mainly from non-ethanol producers) that knowing the production cost of sugar from 

lignocellulosic biomass is important to them as it can help them compare feedstocks and 

pretreatment processes, at least to the extent that these things can be considered modular to their 

process. To that end, NREL maintains a process model for intermediate sugar production that 

reports a minimum sugar selling price (MSSP, $/lb) [86, 87]. It should be stressed that the sugar 

stream produced in this analysis is strictly “imaginary.” The purpose of this analysis is merely to 

separate the cost of producing sugars from the downstream costs of producing ethanol or other 

products.

In the sugar cost analysis, enzymatic hydrolysis is assumed to be carried out in a sterile way and 

nearly to completion so that a transferable sugar stream is produced. The sugar model basically 

follows the ethanol model through enzymatic hydrolysis. A lignin press with counter-current 

washing is added after hydrolysis to separate lignin and unreacted insoluble solids from the 

dilute mixed sugar stream. The combustion section is retained, and an electricity co-product 

credit is assumed for these solids. A wastewater cost is kept to account for treatment of the 

pretreatment flash vapor; however, given that the beer column stillage (the principle wastewater 

feed in the ethanol model) is not applicable in the sugar model, an on-site wastewater facility is 

no longer justified. Instead, an operating cost of $0.09/kg COD is applied for disposal of the 

wastewater material to off-site treatment. Fermentation, distillation, stillage treatment, and 

ethanol storage are completely removed from the sugar model. As in the ethanol process design, 

enzymatic hydrolysis is assumed to achieve 90% cellulose-to-glucose conversion. 

Mass balances and capital and operating costs were determined for the sugar process using the 

same DCFROR methodology described for the ethanol process in Section 4, and the minimum 

dilute sugar selling price was calculated to be 11.6 cents/lb (2007$); detailed economics for this 

case are shown in Figure 19. Table 39 shows the sugar stream composition. It is important to 

note that this value is not directly comparable to a traditional commodity sugar derived from 

corn or sugarcane, as the cellulosic sugar material contains a mixture of C6 and C5 sugars at a 
77 

relatively dilute concentration (set by solids loading limitations in enzymatic hydrolysis) as well 

as other contaminants such as furfurals, inorganic salts, residual lignin, etc. 

Table 39. Mixed Sugar Stream Composition 

Component Transfer

Concentration 

(g/L)

Upgraded 

Concentration 

(g/L)

Glucose 75.9 289.8

Xylose 42.1 160.8

Arabinose 5.1 19.6

Mannose 1.3 4.8

Galactose 3.0 11.5

Total sugars 127.4 486.6

Extractive organics 31.2 119.1

Solubilized lignin 1.7 6.4

HMF 0.9 3.6

Furfural 0.9 0.1

Ammonium sulfate 5.5 21.0

Ammonium acetate 4.1 15.8

Insoluble solids 0.3 1.1

Table 39 also presents a composition for an “upgraded” sugar stream. Given the dilute nature of 

the raw sugar stream at 127 g/L total sugars, the associated selling price of 11.6 cents/lb should 

be viewed more as a “transfer price” where the raw cellulosic sugar material is sold over the 

fence to a conversion process that can use dilute sugar. In order to present a sugar price that more 

closely resembles traditional commodity sugars, we also evaluated the cost of concentrating the 

material on-site. A triple-effect evaporator system was added to the model, with heat input 

specified to achieve 50% water in the syrup (less than 50% sugar given the additional 

components present). The heat for concentrating the sugar syrup was assumed to come from 

supplemental natural gas rather than from the residual lignin. The 2007 price for natural gas for 

industrial use was taken from the U.S. Energy Information Administration, at $7.68 per thousand 

cubic feet [88]. After considering the added cost of the evaporator and condenser equipment as 

well as the cost of natural gas, the minimum sugar selling price was calculated to be 14.3 cents/lb 

(2007$), associated with 487 g/L total sugars. This concentrated sugar still has non-sugar 

compounds that may require cleanup, at additional cost. 
78 

Figure 19. Economic summary for dilute sugar production

Minimum Sugar Selling Price $0.1158 /lb, Dilute sugars

Sugar Production (MMlb / Year) 910.4

Sugar Yield (lb / Dry U.S. Ton Feedstock) 1179

Feedstock Cost $/Dry U.S. Ton $58.5 (Includes handling)

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 40%

Capital Costs

Pretreatment & Conditioning $33,000,000 Feedstock + Handling 4.963

Enzymatic Hydrolysis $19,600,000 Sulfuric Acid 0.164

On-Site Enzyme Production $18,300,000 Ammonia 0.483

Solids Recovery $7,300,000 Glucose (Enzyme Production) 1.296

Wastewater Treatment $0 Other Raw Materials 0.107

Storage $1,900,000 Waste Disposal 0.222

Boiler/Turbogenerator $66,000,000 Electricity -0.676

Utilities $6,900,000 Natural Gas (sugar concentration) 0.000

Total Installed Equipment Cost $152,900,000 Fixed Costs 0.979

Capital Depreciation 0.978

Added Costs $127,000,000 Average Income Tax 0.559

 (% of TPI) 45% Average Return on Investment 2.505

Total Project Investment $279,900,000 Manufacturing Costs ($/yr)

Feedstock + Handling $45,200,000

Installed Equipment Cost/Annual lb $0.17 Sulfuric Acid $1,500,000

Total Project Investment/Annual lb $0.31 Ammonia $4,400,000

Glucose (Enzyme Production) $11,800,000

Loan Rate 8.0% Other Raw Matl. Costs $1,000,000

Term (years) 10 Waste Disposal $2,000,000

Capital Charge Factor 0.131 Electricity -$6,100,000

Natural Gas (sugar concentration) $0

Fixed Costs $8,900,000

Sugar Concentration (g/L) 127 Capital Depreciation $8,900,000

Energy Efficiency (LHV sugar/LHV feed) 50.8% Average Income Tax $5,100,000

Average Return on Investment $22,800,000

Excess Electricity (KWH/lb) 0.12

Plant Electricity Use (KWH/lb) 0.26

Plant Water Use (gal/lb) 0.4

Specific Operating Conditions

Enzyme Loading (mg/g cellulose) 20

Saccharification Time (days) 3.5

Manufacturing Costs (cents/lb sugar)

Sugar Production Process Engineering Analysis

Corn Stover Design Report Case: DW1102A

Dilute Acid Prehydrolysis with Enzymatic Saccharification

All Values in 2007$
79 

5.4 Cost Sensitivity Analysis

Here we present a limited sensitivity analysis using the technoeconomic model presented in this 

report. More exhaustive sensitivity analyses are planned for future publication, including 

sensitivity to feedstock variability and an overarching Monte Carlo sensitivity analysis including 

conversion and economic assumptions. 

A single-point sensitivity was performed on the Aspen model using the variables and limits 

shown in Table 40. The baseline for all variables was the 2012 performance target used in the 

design case described previously. Reasonable minima and maxima for each variable were chosen 

with a group of NREL researchers. Each variable was changed to its maximum and minimum 

value while all others were held constant. The sensitivities of MESP and ethanol yield to these 

parameters are displayed as tornado charts in Figure 20. These indicate that conversion 

parameters have the largest (and only) impacts on ethanol yield. Conversion parameters also 

factor heavily into MESP, along with parameters that affect capital and operating costs. The 

uncertainty in capital costs resulting from the factored approach we used (+/-25%) has the largest 

impact to MESP but is on the order of other key parameters. We note that this is an imperfect 

comparison because the TCI bar represents a true uncertainty while the other bars are more 

accurately called a risk, in the sense that they represent a range of results that we might expect to 

see in the pilot-scale demonstration runs. The magnitude of these bars is of course a function of 

the range we chose in Table 40. Conversion parameters have their own uncertainty due to the 

laboratory measurements involved. We plan to quantify the sensitivity of MESP to these 

uncertainties in a future paper.

Table 40. Assumptions Varied in the Sensitivity Analysis

Assumption Name Min Baseline Max

Pretreatment PT % solids 25% 30% 40%

PT acid loading (mg/g) 10 22.1 35

PT temperature °C 150 158 170

PT xylan to xylose 80% 90% 92%

PT xylan to furfural 3% 5% 8%

PT glucan to glucose 6% 9.9% 12%

Enz hydrolysis EH % solids 17.5% 20% 25%

EH cellulose to glucose 75% 90% 95%

EH enzyme loading mg/g 10 20 30

EH time (d) 2 3.5 5

Fermentation FERM contamination losses 0% 3% 6%

FERM xylose to ethanol 75% 85% 90%

FERM arabinose to ethanol 0% 85% 85%

FERM time (d) 1 1.5 3

Capital Total capital investment -25% - +25%
80 

Figure 20. Single-point sensitivity tornado charts for (a) MESP and (b) ethanol yield

$(0.30) $(0.20) $(0.10) $- $0.10 $0.20 $0.30 

PT xylan to furfural 3:5:8%

PT temperature 150:158:170C

PT % solids 40:30:25%

PT glucan to glucose 12:10:6%

FERM time 1:1.5:3 d

EH time 2:3.5:5 d

FERM arabinose to ethanol 85:85:0%

PT xylan to xylose 92:90:80%

FERM xylose to ethanol 90:85:75%

FERM contamination losses 0:3:6%

EH % solids 25:20:17.5%

PT acid loading 10:22:35 mg/g

EH cellulose to glucose 95:90:75%

EH enzyme loading 10:20:30 mg/g

TCI +/-25%

ΔMESP ($/gal)

Base case $2.15 (a)

-8.0 -6.0 -4.0 -2.0 0.0 2.0 4.0

EH time 2:3.5:5 d

FERM time 1:1.5:3 d

PT % solids 40:30:25%

PT xylan to furfural 3:5:8%

EH enzyme loading 10:20:30 mg/g

PT temperature 150:158:170C

PT acid loading 10:22:35 mg/g

EH % solids 25:20:17.5%

PT glucan to glucose 12:10:6%

FERM arabinose to ethanol 85:85:0%

PT xylan to xylose 92:90:80%

FERM xylose to ethanol 90:85:75%

FERM contamination losses 0:3:6%

EH cellulose to glucose 95:90:75%

ΔYield (gal/ton)

Base case 79.0 (b)
81 

5.5 State of Technology Back-Casting

The annual State of Technology (SOT) assessment is an essential activity for biochemical 

conversion research at NREL. The state of technology process is complementary to the design 

report. Without reconfiguring the unit operations in the Aspen model, the projected conversions 

from the design report are replaced by present conversions that have been experimentally 

verified at the bench and/or pilot scale on process-relevant material. By comparing the SOT 

year-over-year, research advances can be quantified in terms of economic improvements in the 

overall process. 

We understand that “State of Technology” is arguably a misnomer because no commercial 

cellulosic ethanol plants exist today. The aim of an SOT case for a given year is not to report a 

current commercial production cost for cellulosic ethanol, but to reflect NREL’s best estimate of 

ethanol production costs in a hypothetical n

th plant using the current best slate of demonstrated 

technical capabilities. The SOT uses conversion performances that have been demonstrated 

experimentally, with a preference for data obtained in NREL’s ~1 ton/day pilot facility. Not all 

conversion parameters can be practically measured in the pilot plant, so the SOT cases contain 

both laboratory-scale data and pilot-scale data. There is of course some risk in assuming that 

bench-scale performance data is applicable at large scale. 

Since its inception, the SOT analysis has been based on the technoeconomic model from the 

2002 design report [2]. Feedstock costs and year-dollars were updated, and some minor process 

modifications were made, but the basic Aspen model stayed the same. Using the new Aspen 

model described in this report, we performed a back-casting of the 2007–2010 SOTs using their 

conversion results. Table 41 shows these back-cast cases and their related MESP values. Much 

of the research that went into demonstrating these cases was discussed in Sections 3.2 and 3.3.

Also shown in Table 41 are the 2011 interim conversion targets. 
82 

Table 41. Back-Casted State of Technology Cases and Future Targets 

2007 

SOT

2008 

SOT

2009 

SOT

2010

SOT

2011 2012

Minimum Ethanol Selling Price ($/gal) $3.53 $3.46 $3.08 $2.67 $2.62 $2.15

Feedstock contribution $1.01 $0.94 $0.85 $0.72 $0.76 $0.74

Enzyme contribution $0.39 $0.38 $0.36 $0.36 $0.43 $0.34

Non-enzyme conversion contribution $2.14 $2.14 $1.87 $1.59 $1.43 $1.07

TCI/annual gallon $11.33 $11.32 $10.60 $10.15 $9.40 $6.92

Yield (gallon/dry ton) 69 70 76 75 78 79

Technical targets

Feedstock

Feedstock cost ($/dry ton) $69.60 $65.30 $62.05 $53.70 $59.60 $58.50

Pretreatment

Xylan to xylose (including enzymatic) 75% 75% 84% 85% 88% 90%

Xylan to degradation products 13% 11% 6% 8% 5% 5%

Hydrolysate solid-liquid separation Yes Yes Yes Yes Yes No

Xylose sugar loss 2% 2% 2% 2% 1% 1%

Glucose sugar loss 1% 1% 1% 1% 1% 0%

Enzymatic hydrolysis & fermentation 

Enzyme loading (mg/g) 20 20 20 20 25 20

Combined sacch & ferm time (d) 7 7 7 5 5 5

Corn steep liquor loading (wt %) 1% 1% 1% 1% 0.6% 0.25%

Overall cellulose to ethanol 86% 86% 84% 86% 86% 86%

Xylose to ethanol 76% 80% 82% 79% 85% 85%

Arabinose to ethanol 0% 0% 51% 68% 80% 85%

Operating parameters

Pretreatment solids loading (wt %) 30% 30% 30% 30% 30% 30%

Pretreatment temperature (°C) 190 190 158 158 158 158

Acid loading (mg/g dry biomass) 38.0 30.0 24.5 22.1 22.1 22.1

Secondary oligomer hold step No Yes Yes Yes Yes Yes

Ammonia loading (g per L of liquor) 12.9 12.9 9.8 4.8 4.8 4.8

Conditioning mode Liquor Liquor Liquor Liquor Liquor Whole 

slurry

Saccharification mode Washed 

solids

Washed 

solids

Washed 

solids

Washed 

solids

Whole 

slurry

Whole 

slurry

Enz. hydrolysis solids loading (wt %) 20% 20% 20% 17.5% 20% 20%

Reference Aspen model (DW-) 1102F 1102E 1102D 1102C 1102B 1102A

The enzyme costs in Table 41 were computed by the on-site enzyme model described in Section 

3.4. The loadings were set for each case, and the model computed the amount of enzyme needed 

and its cost. No research advances were assumed in the enzyme production section, so in effect 

the unit cost of enzyme protein ($/kg) remained constant across all cases. The enzyme loading 

itself is flat at 20 mg enzyme protein/g cellulose from 2007–2010 and then increases to 25 mg/g 

in 2011. This reflects a change in process integration from washed-solids to whole-slurry 

enzymatic hydrolysis. The rationale for this change was that it will allow us to eliminate all 

solid-liquid separation steps in the process once a suitable whole-slurry hydrolysate conditioning 

process is found. To date, whole-slurry experiments have indicated that a higher enzyme loading 

is required to reach acceptable conversion yields of cellulose to glucose. Reducing this enzyme 

loading back to 20 mg/g is a research target for 2012. 
83 

6 Concluding Remarks

6.1 Summary

NREL analysts and process engineers made a significant effort in 2009–2010 to update the 2002 

design report of Aden et al. [2] that described the biochemical conversion of lignocellulosic 

biomass (corn stover) to ethanol. The purpose of these design reports is to investigate the process 

design and economics of converting lignocellulosic biomass to ethanol at scale and at a profit. 

The end result of the analysis is an absolute selling price for ethanol that is based on reasonable 

process engineering and plant design assumptions. This information is used to predict the costcompetitiveness and market penetration potential of cellulosic ethanol in comparison with 

petroleum-derived fuels and corn ethanol. The design report also helps NREL and DOE to direct 

their research by using it to examine the sensitivity of the selling price to process alternatives, 

equipment and operating cost reductions, and advances in conversion research. Proposed 

research and its anticipated results can be translated into a new design with new economics that 

can be compared with the benchmark case.

In updating the design report, NREL and Harris Group revisited all major assumptions of the 

2002 and earlier design reports, questioning their validity and making adjustments or 

improvements where needed. This included taking advantage of new technologies and 

engineering practices. All equipment costs and major raw material costs were updated with new 

quotes obtained in 2009$ and 2010$. 

Some of the notable conversion process changes from the 2002 design were the following: 

• The corn stover feedstock composition used in the design was updated to reflect a 

composition with a carbohydrate profile closer to the mean observed in NREL studies. 

• A highly detailed pretreatment reactor configuration was provided by Andritz, Inc. The 

specifications of the reactor were mutually agreed upon between Andritz, Harris Group, 

and NREL researchers. 

• Ammonia conditioning of the pretreated biomass replaced the previous practice of 

overliming followed by gypsum precipitation. The ammonia process allows the 

hydrolysate to be conditioned as a whole slurry, eliminating a solid-liquid separation step.

• The enzymatic hydrolysis and fermentation section was previously modeled as a 

continuous train but is now modeled as a batch system. Upstream of the batch reactors, a 

continuous high-solids enzymatic hydrolysis reactor was added. 

• An on-site enzyme production section replaced the previous purchased-enzyme cost 

model. Cellulase enzyme is produced using a T. reesei analog on glucose substrate. The 

purpose of the enzyme production section was to improve transparency on the true cost of 

enzymes for cellulosic ethanol and not to make a judgment call as to whether enzymes 

should be made on-site or purchased from a distributor.

• The wastewater treatment section was completely redesigned in high detail to accurately 

model treatment of the elevated nitrogen and sulfur levels that resulted from the switch 

from overliming to ammonia conditioning. 
84 

The end result of the technoeconomic analysis was a predicted minimum ethanol selling price 

(MESP) of $2.15/gal (2007$). This reflects a $0.74/gal contribution from feedstock, a $0.34/gal 

contribution from enzymes, and a $1.07/gal contribution from the remainder of the conversion 

process.

6.2 Variance from the 2002 Design

In the original 2002 design report, which projected 2012 technology, the modeled 2012 MESP 

was $1.07/gal ethanol in 2000$ ($1.29/gal if directly converted to 2007$). Since then, the 

projected 2012 model has evolved to include inflation in capital and chemical costs, new 

feedstock costs from INL, elimination of feedstock handling and yard workers, and some key 

process changes, including the oligomer conversion step in pretreatment and the switch to 

ammonia conditioning. The most recently updated MESP target for 2012 was $1.49/gal ethanol 

in 2007$ [27]. The model used to compute this price was based on the 2008 State of Technology 

model [21], which in turn was based on the 2002 design model. Table 42 compares this previous 

model to the present design, with the costs broken down by process area (note that the totals may 

be appear to be off by a cent due to rounding). 

Table 42. Comparison of the Previous Target Case and the Present Design (2007$/gal) 

Model Comparison by Area 2012 Model Based 

on 2008 SOT

New 2012 Model 

(This Work) 

Difference

Feedstock $0.57 $0.74 $0.17

Feed handling (included in 

feedstock cost in current model) $0.01 $0.00 -$0.01

Pretreatment/conditioning $0.26 $0.29 $0.03

Saccharification and fermentation $0.12 $0.20 $0.08

Cellulase $0.12 $0.34 $0.22

Distillation and solids recovery $0.16 $0.12 -$0.04

Wastewater treatment $0.03 $0.34 $0.31

Storage $0.02 $0.02 $0.01

Boiler/turbogenerator (includes 

electricity co-product credit) $0.16 $0.04 -$0.12

Utilities $0.06 $0.06 $0.01

TOTAL (MESP) $1.49 $2.15 $0.66

Because the MESP is expressed in $/gal, it depends of course on the overall ethanol yield, which 

in the present design is about 12% lower than in the older model. Overall, there was a $0.66/gal 

(44%) increase over the previous target of $1.49/gal. When broken down by process area, we see 

in Table 42 that the more notable changes in cost came from the enzyme section (due to a change 

in the way enzyme costs were determined, i.e., by on-site production instead of an assumed price 

per gallon of ethanol) and from the wastewater area (material that was previously burned is now 

being routed to wastewater). The individual contributors to the cost delta are expressed in a 

different way in Table 43. At $0.22 per gallon, the new enzyme cost model is the largest 

contributor to the MESP difference between the two models. Another $0.08 can be attributed to 

the lower yield from the new feedstock composition, and $0.10 to an increase in the feedstock 

cost. A total of $0.21 can be attributed to net increases in the capital equipment costs, chemical 

costs, and labor rates.
85 

Table 43. Individual Contributors to the Change in MESP

Contributor Amount

(2007$/gal)

Explanation

Feedstock cost $0.10 From $50.90 to $58.50 (new grower payment assumption)

Feedstock composition $0.08 Lower yield from more representative feedstock composition

Cellulase $0.22 On-site production model predicted $0.34/gal versus the 

previous $0.12/gal purchased model

Electricity credit -$0.07 Larger fraction of biomass is converted to electricity with the 

updated composition

Capital $0.30 Net increase in the capital equipment required

Financial assumptions -$0.06 100% equity versus 40% equity with 8% interest

Chemicals $0.06 New chemical costs

Fixed costs $0.03 New labor costs

$0.66

 

6.3 Future Work

The overarching near-term goal for DOE’s Office of the Biomass Program is to demonstrate the 

cost-competitiveness of cellulosic ethanol with petroleum fuels by 2012. NREL and DOE have 

agreed on a slate of conversion parameters that are likely achievable by 2012. These 

development targets lead to an n

th-plant MESP of $2.15/gal by 2012, as modeled by the process 

design described in this report. One of NREL’s main tasks is to perform integrated pilot-scale 

testing that demonstrates these conversion targets in 2012. The pilot-scale tests will make use of 

NREL’s biochemical process integration expertise and other core conversion research occurring 

at smaller scales. 

Pretreatment research will continue to focus on increasing xylan-to-xylose conversion, both by 

optimizing the conditions between the first step of pretreatment and the secondary oligomer 

conversion step and by supplementing enzymatic hydrolysis with accessory xylanase enzymes. 

Enzymatic hydrolysis research will have a significant focus on integrated experiments and on 

optimizing conditions using the newest pretreated material with the highest xylose conversion. 

Another focus in the pretreatment and enzymatic hydrolysis areas will be to reduce chemical 

usage and severity in pretreatment while maintaining good enzyme digestibility of the pretreated 

biomass. These experiments must be performed in an integrated way with fermentation. Also, 

some of the results described in this design report use an advanced research enzyme preparation. 

It is likely we will continue to experiment with more advanced preparations in 2011 and beyond. 

In fermentation research, some of the advanced strains being developed by the awardees of 

DOE’s ethanologen strain solicitation are slated to be tested in NREL’s pilot plant. There is also 

a significant internal effort to genetically engineer the NREL Z. mobilis strain with a metabolic 

pathway to convert arabinose to ethanol. This is a risk mitigation strategy should the solicitation 

strains prove difficult to integrate into NREL’s pilot-scale ethanol process. 

Process integration research will continue to focus on cost reduction through process design. 

Obtaining a pretreatment, conditioning, and enzymatic hydrolysis process that can be carried out 

as a whole slurry, without a solid-liquid separation step, is key to future cost reductions and will 

be studied aggressively. Other areas where process integration research plays a role is in 

reducing costs of fermentation media and mitigating the problems in back-end cleanup and 
86 

wastewater treatment discussed in this report. These areas of the process will receive much more 

scrutiny between now and 2012. 

In technoeconomic analysis research, we have plans to perform an extensive Monte Carlo costsensitivity analysis using the new design report process described here. The analysis will 

examine the sensitivity to economic parameters such as raw material and equipment costs, to the 

expected compositional variability of corn stover, and to the specific process parameters defined 

in Aspen, such as fermentation conversions. There is also an interest in updating the Aspen 

model to handle alternate feedstocks: switchgrass, sorghum, wheat straw, sugarcane bagasse, etc. 

Each feedstock may require key process changes to overcome biomass recalcitrance or a thermal 

deficit. Once we have determined the similarities and difference between feedstocks, the goal is 

to produce a single, flexible model that can handle them all.

In addition to alternate feedstocks, this work is applicable to further technoeconomic analysis of 

advanced biofuels. The cellulosic ethanol process described here is only one of many pathways 

by which a transportation fuel can be produced from biomass. In the future, advanced conversion 

technologies that produce infrastructure-compatible or “drop-in” fuels (primarily gasoline, 

diesel, and/or jet-range hydrocarbons) will have an increasing importance in NREL’s and DOE’s 

research portfolio. Figure 21 shows some example pathways that yield such fuels from sugar; 

most of these technologies are already being researched by startup companies and academic 

institutions using clean sugar from corn or sugar cane. NREL plans to leverage its expertise in 

sugar generation from lignocellulosic material to help move these technologies away from foodbased sugars and toward sustainable biomass.

Figure 21. Sugar-based pathways for conventional and advanced biofuels

In a technoeconomic model of an advanced conversion process using cellulosic sugar, the front 

end of the Aspen model described in this report would largely stay the same through 

pretreatment and enzymatic hydrolysis, up to the production of a sugar intermediate. The sugar 

conversion and product recovery sections would be replaced with new sections that describe the 

technologies in Figure 21. Our process models for sugar production and the specialized 

wastewater and utilities sections should therefore be considered modular to these future studies. 
87 

References 

NREL milestone reports cited below cannot be accessed outside of NREL and DOE. Readers 

may contact the authors of the specific reference to determine if this information has been made 

public since publication of this design report. 

1. Wooley, R.; Ruth, M.; Sheehan, J.; Ibsen, K. Lignocellulosic Biomass to Ethanol Process 

Design and Economics Utilizing Co-Current Dilute Acid Prehydrolysis and Enzymatic 

Hydrolysis: Current and Future Scenarios. Report No. NREL/TP-580-26157. Golden, 

CO: National Renewable Energy Laboratory, July 1999. 

http://www.nrel.gov/docs/fy99osti/26157.pdf. 

2. Aden, A.; Ruth, M.; Ibsen, K.; Jechura, J.; Neeves, K.; Sheehan, J.; Wallace, B.; 

Montague, L.; Slayton, A.; Lukas, J. Lignocellulosic Biomass to Ethanol Process Design 

and Economics Utilizing Co-Current Dilute Acid Prehydrolysis and Enzymatic 

Hydrolysis for Corn Stover. Report No. NREL/TP-510-32438. Golden, CO: National 

Renewable Energy Laboratory, June 2002. http://www.nrel.gov/docs/fy02osti/32438.pdf. 

3. “Aspen Plus.” Burlington, MA: Aspen Technology. 

http://www.aspentech.com/products/aspen-plus.cfm. 

4. Tao, L.; Aden, A. “The Economics of Current and Future Biofuels.” In Vitro Cell Dev. 

Biol.; Vol. 45, 2009; pp. 199-217. 

5. F.O. Lichts. Ethanol Production Costs: A Worldwide Survey. Turnbridge Wells, Kent, 

UK: Agra Informa, 2007. 

6. Kwiatkowski, J.R.; McAloon, A.J.; Taylor, F. “Modeling the Process and Costs of Fuel 

Ethanol Production by the Corn Dry-Grind Process.” Ind. Crops Prod.; Vol. 233, 2006; 

pp. 288-296. 

7. Rodrigues, A.P. Participação dos fornecedores de cana na cadeia do açúcar e álcool. 

Congresso Internacional de Tecnologias na Cadeia Produtiva, Concana Uberaba, MG, 

março de, 2007 (in Portuguese). 

8. Seabra, J.E.A. Technical-Economic Evaluation of Options for Whole Use of Sugar Cane 

Biomass in Brazil. Ph.D. Thesis. Campinas Faculdade de Engenharia Mecânica 

Universidade Estadual de Campinas, 2007 (in Portuguese). 

9. Graboski, M.S.; McCormick, R.L. “Combustion of Fat and Vegetable Oil Derived Fuels 

in Diesel Engines.” Prog. Energy Combust. Sci.; Vol. 24, 1998; pp. 125–164. 

10. Haas, M.J.; McAloon, A.J.; Yee, W.C.; Foglia, A.T. “A Process Model to Estimate 

Biodiesel Production Cost.” Biores. Technol.; Vol. 97, 2006; pp. 671–678. 

11. Laser, M.; Larson, E.; Dale, B.; Wang, M.; Greene, N.; Lynd, L.R. “Comparative 

Analysis of Efficiency, Environmental Impact, and Process Economics for Mature 
88 

Biomass Refining Scenarios.” Biofuels, Bioproducts, & Biorefining; Vol. 3, 2009; pp. 

247-270. 

12. Kazi, F.K.; Fortman, J.A.; Anex, R.P.; Hsu, D.D.; Aden, A.; Dutta, A.; Kothandaraman, 

G. “Techno-Economic Comparison of Process Technologies for Biochemical Ethanol 

Production from Corn Stover.” Fuel; Vol. 89, 2010; pp. S20-S28. 

13. Klein-Marcuschamer, D.; Oleskowicz-Popiel, P.; Simmons, B.A.; Blanch, H.W. 

“Technoeconomic Analysis of Biofuels: A Wiki-Based Platform for Lignocellulosic 

Biorefineries.” Biomass and Bioenergy; Vol. 34, 2010; pp. 1914-1921. 

14. Klein-Marcuschamer, D.; Oleskowicz-Popiel, P.; Simmons, B.A.; Blanch, H.W. “JBEI 

Corn Stover to Ethanol Model.” U.S. Department of Energy Joint BioEnergy Institute 

Wiki, http://econ.jbei.org/wiki/index.php/Corn_Stover_to_Ethanol_Model. 

15. Huang, H.J.; Ramaswamy, S.; Al-Dajani, W.; Tschirner, U.; Cairncross, R.A. “Effect of 

Biomass Species and Plant Size on Cellulosic Ethanol: A Comparative Process and 

Economic Analysis.” Biomass and Bioenergy; Vol. 33, 2009; pp. 234-246. 

16. Sendich, E.N.; Laser, M.; Kim, S.; Alizadeh, H.; Laureano-Perez, L.; Dale, B.; Lynd, L. 

“Recent Process Improvements for the Ammonia Fiber Expansion (AFEX) Process and 

Resulting Reductions in Minimum Ethanol Selling Price.” Bioresource Technology; Vol. 

99, 2008; pp. 8429-8435. 

17. Gnansounou, E.; Dauriat, A. “Techno-Economic Analysis of Lignocellulosic Ethanol: A 

Review.” Bioresource Technology; Vol. 101, 2010; pp. 4980-4991. 

18. Bals, B.; Wedding, C.; Balan, V.; Sendich, E.; Dale, B. “Evaluating the Impact of 

Ammonia Fiber Expansion (AFEX) Pretreatment Conditions on the Cost of Ethanol 

Production.” Bioresource Technology; Vol. 102, 2011; pp. 1277-1283. 

19. Piccolo, C.; Bezzo, F. “A Techno-Economic Comparison Between Two Technologies for 

Bioethanol Production from Lignocellulose.” Biomass and Bioenergy; Vol. 33, 2009; pp. 

478-491. 

20. National Academy of Sciences. Liquid Transportation Fuels from Coal and Biomass: 

Technological Status, Costs, and Environmental Impacts. Washington, DC: The National 

Academies Press, 2009. http://www.nap.edu/catalog/12620.html. 

21. Humbird, D.; Aden, A. Biochemical Production of Ethanol from Corn Stover: 2008 State 

of Technology Model. Report No. NREL/TP-510-46214. Golden, CO: National 

Renewable Energy Laboratory, July 2009. http://www.nrel.gov/docs/fy09osti/46214.pdf. 

22. Wooley, R; Putsche, V. Development of an ASPEN PLUS Physical Property Database 

for Biofuels Components. Report No. NREL/MP-425-20685. Golden, CO: National 

Rnewable Energy Laboratory, April 1996. http://biodev.nrel.gov/pdfs/3955.pdf. 
89 

23. Wooley, R.; Ibsen, K.; Putsche, V. Development of an ASPEN PLUS Physical Property 

Database for Biofuels Components. NREL Technical Memo, Document #4141. Golden, 

CO: National Renewable Energy Laboratory, 1999. 

http://biodev.nrel.gov/bcfcdoc/4141.pdf. 

24. Hess, J.R.; Kenney, K.L.; Ovard, L.P.; Searcy, E.M.; Wright, C.T. Uniform-Format Solid 

Feedstock Supply System: A Commodity-Scale Design to Produce an InfrastructureCompatible Bulk Solid from Lignocellulosic Biomass, Section 3: Pioneer Uniform 

Feedstock Supply System. Report No. INL/EXT-08-14752. Idaho Falls, ID: Idaho 

National Laboratory, 2009. http://www.inl.gov/bioenergy/uniform-feedstock. 

25. Templeton, D.W.; Sluiter, A.D.; Hayward, T.K.; Hames, B.R.; Thomas, S.R. “Assessing 

Corn Stover Composition and Sources of Variability via NIRS.” Cellulose; Vol. 16, 

2009; pp. 621-639. 

26. Corn stover analysis, sample 33A14 from Kramer Farm, Wray Pioneer Lot. Huffman 

Report #123409. Wray, CO: Kramer Farm. 

27. Office of the Biomass Program, Energy Efficiency and Renewable Energy, U.S. 

Department of Energy. Multi-Year Program Plan. Washington, DC: U.S. Department of 

Energy, April 2011. http://www1.eere.energy.gov/biomass/pdfs/mypp_april_2011.pdf. 

28. Chen, S.F.; Mowery, R.A.; Scarlata, C.J.; Chambliss, C.K. “Compositional Analysis of 

Water-Soluble Materials in Corn Stover.” J. Agric. Food Chem.; Vol. 55, 2007; pp. 5912-

5918. 

29. Bower, S.; Wickramasinghe, R.; Nagle, N.J.; Schell, D.J. “Modeling Sucrose Hydrolysis 

in Dilute Sulfuric Acid Solutions at Pretreatment Conditions for Lignocellulosic 

Biomass.” Biores. Technol.; Vol. 99, 2008; pp. 7354-7362. 

30. Ethanol (data page). Wikipedia, 

http://en.wikipedia.org/wiki/Ethanol_%28data_page%29. 

31. Nagle, N.; Weiss, N.; Kuhn, E.; Tucker, M.; Sievers, D.; Elander, R. Validate > 75% 

Xylan Yield and <8% Degradation from Whole Corn Stover using a High Solids 

Continuous Pretreatment Reactor. D Milestone Report, Document #10341. Golden, CO: 

National Renewable Energy Laboratory, 2008. http://biodev.nrel.gov/bcfcdoc/10341.pdf. 

32. Jennings, E.; Mohagheghi, A.; Schell, D.; Dowe, N. Technical and Economic 

Performance of an Ammonium Hydroxide-Based Conditioning Technology. E Milestone 

Report, Document #10379. Golden, CO: National Renewable Energy Laboratory, 

November 2008. http://biodev.nrel.gov/bcfcdoc/10379.pdf. 

33. Jechura, J. Correlation of Metal Corrosion Data. NREL Technical Memo, Document 

#5194. Golden, CO: National Renewable Energy Laboratory, March 2001. 

http://biodev.nrel.gov/bcfcdoc/5194.pdf. 
90 

34. Humbird, D. FY 2009 State of Technology Update—Biochem. D Milestone, Document 

#10590. Golden, CO: National Renewable Energy Laboratory, October 2009. 

http://biodev.nrel.gov/bcfcdoc/10590.pdf. 

35. Humbird, D. FY 2010 Biochemical Platform State of Technology Update. D Milestone, 

Document #10752. Golden, CO: National Renewable Energy Laboratory, September 

2010. http://biodev.nrel.gov/bcfcdoc/10752.pdf. 

36. Weiss, N.; Nagle, N.; Elander, R.; Tucker, M. Achieve 75% Xylose Yield in Laboratory 

Scale High Solids Pretreatment Reactor on Corn Stem Internode. D Milestone Report, 

Document #9926. Golden, CO: National Renewable Energy Laboratory, 2007. 

http://biodev.nrel.gov/bcfcdoc/9926.pdf. 

37. Sievers, D.; Kuhn, E.; Nagle, N.; Tucker, M.; Weiss, N. Implement Equipment 

Modifications to the Existing Horizontal Pretreatment Reactor System to Allow for 

Operation with a Tighter Residence Time Distribution. D Milestone Report, Document 

#10404. Golden, CO: National Renewable Energy Laboratory, December 2008. 

http://biodev.nrel.gov/bcfcdoc/10404.pdf. 

38. Nagle, N.; Kuhn, E.; Sievers, D.; Tucker, M.; Weiss, N.; Elander, R. Parametric Study

using the Continuous Horizontal Pretreatment Reactor using Corn Stover to Determine 

Operating Conditions that Achieve 80% Yield of Monomeric Xylose. D Milestone Report, 

Document #10575 Golden, CO: National Renewable Energy Laboratory, September 

2009. http://biodev.nrel.gov/bcfcdoc/10575.pdf. 

39. Nagle, N.; Kuhn, E.; Elander, R. Parametric Study Using the Modified Continuous 

Horizontal Pretreatment Reactor Using Corn Stover to Determine Operating Conditions 

that Can Achieve Yields of ≥ 85% of Monomeric Xylose. NREL D Milestone Report, 

Document #10764 Golden, CO: National Renewable Energy Laboratory, September 

2010. http://biodev.nrel.gov/bcfcdoc/10764.pdf. 

40. Selig, M; Decker, S. Enzymatic Conversion of Oligosaccharides from Various 

Pretreatment Hydrolysates. Joule Milestone Report, Document #10114. Golden, CO: 

National Renewable Energy Laboratory, March 2008. 

http://biodev.nrel.gov/bcfcdoc/10114.pdf. 

41. Humbird, D.; Elander, R. Develop Process Engineering Scenarios for Potential Leading 

Pretreatment/Enzymatic Hydrolysis Configurations. E Milestone Report, Document 

#10582. Golden, CO: National Renewable Energy Laboratory, September 2009. 

http://biodev.nrel.gov/bcfcdoc/10582.pdf. 

42. Zhang, M.; Eddy, C.; Deanda, K.; Finkelstein, M.; Picataggio, S. “Metabolic Engineering 

of a Pentose Metabolism Pathway in Zymomonas mobilis.” Science; Vol. 267, 1995; pp. 

240-243. 

43. Miroslav, S; Ho, N. “Production of ethanol from cellulosic biomass hydrolysates using 

genetically engineered saccharomyces yeast capable of cofermenting glucose and 

xylose,” Applied Biochemistry and Biotechnology; Vol. 144, 2004; pp. 403-416. 
91 

44. “Integrated Biorefinery Research Facility.” National Renewable Energy Laboratory 

website, http://www.nrel.gov/biomass/integrated_biorefinery_research_facility.html. 

45. Humbird, D.; Mohagheghi, A.; Dowe, N.; Schell, D. “Economic Impact of Total Solids 

Loading on Enzymatic Hydrolysis of Dilute-Acid Pretreated Corn Stover.” Biotechnology 

Progress; Vol. 26, 2010; pp.1245-1251. 

46. Loos, H.; Kramer, R.; Sahm, H.; Sprenger, G. “Sorbitol Promotes Growth of Zymomonas 

mobilis in Environments with High Concentrations of Sugar: Evidence for a 

Physiological Function of Glucose-Fructose Oxidoreductase in Osmoprotection.” J. 

Bact.; Vol. 179, 1994; pp. 7688-7693. 

47. Dowe, N.; Humbird, D.; Schell, D. Status Report on Overall Corn Stover to Ethanol 

Yield Improvement Effort. Joule Milestone, Document #10457. Golden, CO: National 

Renewable Energy Laboratory, March 2009. http://biodev.nrel.gov/bcfcdoc/10457.pdf. 

48. Franden, M.A.; Zhang, M. Characterize Hydrolyzate Toxicity Across a Variety of 

Hydrolyzates Obtained from the Horizontal Pretreatment Reactor and Other Selected 

Process Relevant Hydrolyzates Using both Biological and Chemical Analysis. D 

Milestone, Document #10585. Golden, CO: National Renewable Energy Laboratory, 

September 2009. http://biodev.nrel.gov/bcfcdoc/10585.pdf. 

49. “Enzyme Systems Solicitation.” Financial Opportunities: Past Solicitations. Washington, 

DC: U.S. Department of Energy, Energy Efficiency and Renewable Energy website, 

2008. http://www1.eere.energy.gov/biomass/past_solicitations.html#Enzymes. 

50. “Ethanologen Projects—Development of Fermentative Microorganisms.” Financial 

Opportunities: Past Solicitations. Washington, DC: U.S. Department of Energy, Energy 

Efficiency and Renewable Energy website, 2007. 

http://www1.eere.energy.gov/biomass/past_solicitations.html#Ethanologen_Projects. 

51. Schell, D.; Humbird, D.; Wolfrum, E.; Zhang, M. Research Plan for Demonstrating the 

2012 Cost Target. E Milestone, Document #10458. Golden, CO: National Renewable 

Energy Laboratory, March 2009. http://biodev.nrel.gov/bcfcdoc/10458.pdf. 

52. “Ethanol Co-Products.” Ethanol Information website, http://ethanolinformation.com/ethanol-coproducts.php. 

53. Meerman, H.J.; Kelley, A.S.; Ward, M. “Advances in Protein Expression in Filamentous 

Fungi,” Chapter 7. Baneyx, F., ed. Protein Expression Technologies: Current Status and 

Future Trends. Norfolk, UK: Horizon Bioscience, 2004; p 345. 

54. Warzywoda, M., et al., U.S. Patent No. 4,762,788, 28 November 1984. 

55. Emme, B., Novozymes, Inc., personal communication with David Humbird, 2009. 

56. Atkinson, B.; Mavituna, F. Biochemical Engineering and Biotechnology Handbook. New 

York: The Nature Press, 1983. 
92 

57. Wang, D.; Cooney, C.; Demain, A.; Dunnill, P.; Humphrey, A.; Lilly, M. Fermentation 

and Enzyme Technology. New York: John Wiley & Sons, 1979. 

58. Schell, D.; Riley, C.; Bergeron, P.; Walter, P. Technical and Economic Analysis of an 

Enzymatic Hydrolysis Based Ethanol Plant – Draft. SERI Technical Report TP-232-

4295. Golden, CO: Solar Energy Research Institute, 1991. 

http://biodev.nrel.gov/bcfcdoc/3892.pdf. 

59. Jechura, J. Sugar Platform Post Enzyme Subcontract Case. Technical Memo, Document 

#9382 Golden, CO: National Renewable Energy Laboratory, October 2005. 

http://biodev.nrel.gov/bcfcdoc/9382.pdf. 

60. Merino, S.T.; Cherry, J. “Progress and Challenges in Enzyme Development for Biomass 

Utilization.” Adv. Biochem. Engin./Biotechnol.; Vol. 108, 2007; pp. 95-120. 

61. Dean, D.; Dodge, T.; Valle, F.; Chotani, G. “Development of biorefineries—technical 

and economic considerations.” Kamm, B. et al., ed. Biorefineries—industrial processes 

and products. Status quo and future directions. Vol. 1, Weinheim, Germany: WileyVCH, 2006; pp. 67–83. 

62. “New Enzymes Turn Waste into Fuel.” Press release. Novozymes, Inc., February 15, 

2010. http://novozymes.com/en/news/news-archive/Pages/45713.aspx. 

63. “Genencor Introduces Accellerase DUET.” Press release. Genencor, February 15, 2010. 

http://www.genencor.com/wps/wcm/connect/genencor/genencor/media_relations/news/ar

chive/2010/pressrelease_509_en.htm. 

64. Ahmetović, E.; Martin, M.; Grossman, I.E. “Optimization of energy and water 

consumption in corn-based ethanol plants,” Ind. Eng. Chem. Res.; Vol. 49, 2010; pp. 

7972-7982. 

65. Steinwinder, T.; Gill, E.; Gerhardt, M. Process Design of Wastewater Treatment for the 

NREL Cellulosic Ethanol Model. NREL/SR-5100-51838. Work performed by Brown and 

Caldwell, Nashville, TN, under NREL Subcontract RGB-0-40607. Golden, CO: National 

Renewable Energy Laboratory, September 2011. 

http://www.nrel.gov/docs/fy11osti/51838.pdf. 

66. Mohagheghi, A.; Ruth, M.; Schell, D. “Conditioning Hemicellulosic Hydrolysates for 

Fermentation: Effects of Overliming pH on Sugar and Ethanol Yields.” Process 

Biochemistry; Vol. 41, 2006; pp. 1806-1811. 

67. Reaction Engineering International. Steam and Electricity Generation Options for the 

Biomass-To-Ethanol Process. Work performed by Reaction Engineering International, 

Salt Lake City, UT, under NREL Subcontract ACO-8-18019-01. Golden, CO: National 

Renewable Energy Laboratory, March 1998. 
93 

68. Philippek, C.; Knöbig, T.; Schönfelder, H.; Werther, J. “NOx Formation and Reduction 

During Combustion of Wet Sewage Sludge in the Circulating Fluidized Bed.” 

Proceedings of the 14th International Conference on Fluidized Bed Combustion; Vol. 2, 

New York: ASME Press, 1997. 

69. Chemical Engineering Magazine Plant Cost Index. Chemical Engineering Magazine. 

http://www.che.com/pci/. 

70. SRI Consulting. “U.S. Producer Price Indexes – Chemicals and Allied Products/Industrial 

Inorganic Chemicals Index.” Chemical Economics Handbook. Menlo Park, CA: SRI 

Consulting, October 2008.

71. “National Employment, Hours, and Earnings Catalog, Industry: Chemicals and Allied 

Products.” Bureau of Labor Statistics Data website, Series ID: CEU3232500008, Years 

1980-2009, annual average. http://data.bls.gov/cgi-bin/srgate. 

72. North American Electric Reliability Corporation. “Wholesale Average Electricity Price 

Table.” U.S. Energy Information Administration website, 

http://www.eia.doe.gov/cneaf/electricity/wholesale/wholesale.html. 

73. Garrett, D.E. Chemical Engineering Economics. New York: Van Nostrand Reinhold, 

1989. 

74. Peters, M.S.; Timmerhaus, K.D. Plant Design and Economics for Chemical Engineers. 

5th Ed., New York: McGraw-Hill, 2003. 

75. Chem Systems. Biomass to Ethanol Process Evaluation. NREL Final Subcontract 

Report. Work performed by Chem Systems, Tarrytown, NY. Golden, CO: National 

Renewable Energy Laboratory, December 1994. http://biodev.nrel.gov/pdfs/3925.pdf. 

76. Walas, S.M. Chemical Process Equipment - Selection and Design. Newton, MA: 

Butterworth Heinemann, 1988.

77. Cran, J. “Improved factored method gives better preliminary cost estimates.” Chemical 

Engineering, April 6, 1981; pp. 65-79. 

78. “Sugar and Sweeteners Yearbook Tables: Excel Spreadsheets.” Table 7—U.S. wholesale 

list price for glucose syrup, Midwest markets, monthly, quarterly, and by calendar and 

fiscal year. United States Department of Agriculture Economic Research Service website, 

http://www.ers.usda.gov/Briefing/Sugar/data.htm#yearbook. Direct download link for 

Table 7: http://www.ers.usda.gov/Briefing/Sugar/Data/TABLE07.XLS. 

79. Urbanchuck, J. Contribution of the Ethanol Industry to the Economy of the United States. 

Prepared by LECG, LLC for the Renewable Fuels Association, 2008. 

80. SRI Consulting. PEP Yearbook. Menlo Park, CA: SRI Consulting, 2008. 

http://www.sriconsulting.com/. 
94 

81. Short, W.; Packey, D.J.; Holt, T. A Manual for the Economic Evaluation and Energy 

Efficiency and Renewable Energy Technologies. Report No. TP-462-5173. Golden, CO: 

National Renewable Energy Laboratory, March 1995. 

82. Internal Revenue Service. How to Depreciate Property. Publication #946. Washington, 

DC: Department of the Treasury Internal Revenue Service, 2009. 

http://www.irs.gov/pub/irs-pdf/p946.pdf. 

83. Perry, R.H.; Green, D.W. Perry’s Chemical Engineers’ Handbook. 7th Ed., New York: 

McGraw-Hill, 1997. 

84. Gary, J.H.; Handwerk, G.E. Petroleum Refining, Technology and Economics. 3rd Ed., 

New York: Marcel Dekker, Inc., 1994. 

85. Macknick, J.; Newmark, R.; Heath, G.; Hallet, K.C. A Review of Operational Water 

Consumption and Withdrawal Factors for Electricity Generating Technologies. Report 

No. NREL/TP-6A20-50900. Golden, CO: National Renewable Energy Laboratory, 

March 2011. http://www.nrel.gov/docs/fy11osti/50900.pdf. 

86. Dutta, A.; Ibsen, K.; Dowe, N. Complete Integrated Tests of Pretreatment and Enzymatic 

Hydrolysis in Conjunction with Existing Fermentation Organisms at Bench-Scale on 

Corn Stover that Validate $0.125 per Pound Sugars on the Pathway to Achieving $0.064 

per Pound in 2012. Joule Milestone Report. Golden, CO: National Renewable Energy 

Laboratory, September 2007. http://devafdc.nrel.gov/bcfcdoc/9985.pdf. Rev. February 

2008. http://devafdc.nrel.gov/bcfcdoc/10087.pdf. 

87. Humbird, D. Demonstrate Alternative Pretreatment Technologies at Bench-Scale using 

Advanced Cellulase Enzymes and Integrated Technologies that have the Potential of 

Achieving $0.12/lb Sugar on the Pathway to $0.073/lb by 2012 (in 2007$). Joule 

Milestone Report, Document #10572. Golden, CO: National Renewable Energy 

Laboratory, September 2009. http://biodev.nrel.gov/bcfcdoc/10572.pdf. 

88. “Natural Gas Prices.” U.S. Energy Information Administration website,

http://tonto.eia.doe.gov/dnav/ng/ng_pri_sum_a_EPG0_PIN_DMcf_a.htm. 

89. Keller and Heckmann LLP. Assessment Plan for Corn Steep Liquor (CAS #66071-94-1) 

in Accordance with the USEPA High Production Volume Chemical Challenge Program. 

Prepared for The Corn Refiners Association, 2006. 

90. Verevkin, S.P.; Emel’yanenko V.N; Stepurko, E.N.; Ralys, R.V,; Zaitsau, D.H.; Stark, A. 

“Biomass-derived platform chemicals: Thermodynamic studies on the conversion of 5-

hydroxymethylfurfural into bulk intermediates.” Ind. Eng. Chem. Res.; Vol. 48, 2009; pp. 

10087-10093. 

91. Lewis Sr., R.J. (2007). Hawley's Condensed Chemical Dictionary. 15th Ed., New York: 

John Wiley & Sons, 2007. 
95 

92. Merrill, A.L.; Watt, B.K. Energy Value of Foods: Basis and Derivation. USDA ARS 

Handbook #74, 1973. 

http://www.ars.usda.gov/SP2UserFiles/Place/12354500/Data/Classics/ah74.pdf. 

93. Roels, J.A. Energetics and Kinetics in Biotechnology. Amsterdam: Elsevier Biomedical 

Press, 1983. 

94. Bailey, J.E.; Ollis, D.F. Biochemical Engineering Fundamentals. 2nd Ed., New York:

McGraw-Hill, 1986. 
96 

Appendix A. Individual Equipment Costs Summary

The left side of the following table shows abbreviated specifications and the purchased cost for 

each piece of equipment in the PFDs. The right side of the table shows the calculations used to 

scale the purchased cost for size and installation. Although each piece of equipment has its own 

line, many were quoted as part of a package so their scaling calculations are grayed out. Note 

that the package prices are generally listed on the first associated line-item. 

NREL and Harris Group would like to acknowledge the equipment vendors who assisted us with 

the cost estimation effort for this design report. 
97 

Mechanical Equipment List Scaled Installed Costs

EQUIPMENT TITLE VENDOR DESCRIPTION HP MATERIAL

NUM REQD

$

Year of 

Quote

Scaling Variable

Scaling Val

Units

Scaling Exp

Inst Factor

New Val

Size Ratio

Scaled Purch 

Cost

Purch Cost in 

Proj year

Inst Cost in Proj 

year

C- 101 Transfer Conveyor Dearborn Midwest 160 MTPH ea., enclosed, 60 in. x 65 ft. long 20 hp ea. CS 2 $5,397,000 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $5,714,628 $5,752,952 $9,780,018

C- 102 High Angle Transfer Conveyor Dearborn Midwest 160 MTPH ea., enclosed, 72 in. wide 50 hp ea. CS 2 INCLUDED

C- 103 Reversing Load-in Conveyor Dearborn Midwest 320 MT / hr, enclosed, 84 in. wide 20 hp CS 1 INCLUDED

C- 104 Dome Reclaim System Cambelt 100 MTPH ea. 45 kw ea. CS 2 $3,046,000 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $3,225,265 $3,246,895 $5,519,721

C- 105 Reclaim Conveyor Dearborn Midwest 100 MTPH ea., enclosed, 48 in. x 125 ft. long 10 hp ea. CS 2 INCLUDED

C- 106 High Angle Transfer Conveyor Dearborn Midwest 100 MTPH, enclosed, 72 in. wide 20 hp CS 1 INCLUDED

C- 107 Elevated Transfer Conveyor Dearborn Midwest 100 MTPH, enclosed, 48 in. x 200 ft. long 10 hp CS 1 INCLUDED

C- 108 Process Feed Conveyor Dearborn Midwest 70 MTPH ea., enclosed, 42 in. x 25 ft. long 5 hp ea. CS 1 INCLUDED

M-101 Truck Scale St. Louis Scale 10' x 70', 200,000 lb CONCRETE 2 $110,000 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $116,474 $117,255 $199,333

M-102 Truck Dumper Jeffrey Rader 70' x 55 ton x 63 degree 2 x 50 hp CS 2 $484,000 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $512,485 $515,922 $877,067

M-103 Truck Dumper Hopper Jeffrey Rader 3500 cu.ft. hopper w/ drag chain conveyor 50 hp ea. CS 2 $502,000 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $531,544 $535,109 $909,685

M-104 Concrete Feedstock Storage Dome Domtec 98 ft. dia., 160 ft. high., 4000 MT CONCRETE 2 $3,500,000 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $3,705,984 $3,730,838 $6,342,424

M-105 Belt Scale Tecweigh Scale plus processor CS 2 $10,790 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $11,425 $11,502 $19,553

M-106 Dust Collection System Sly 8500 ACFM 25 hp ea. CS 6 $279,900 2009 STRM.101 94697 kg/hr 0.60 1.7 104167 1.10 $296,373 $298,360 $507,213

Area 100 Totals $14,114,178 $14,208,831 $24,155,013

EQPT NO
98 

Mechanical Equipment List Scaled Installed Costs

EQUIPMENT TITLE VENDOR DESCRIPTION HP MATERIAL

NUM REQD

$

Year of 

Quote

Scaling Variable

Scaling Val

Units

Scaling Exp

Inst Factor

New Val

Size Ratio

Scaled Purch 

Cost

Purch Cost in 

Proj year

Inst Cost in Proj 

year

A- 201 In-line Sulfuric Acid Mixer KOMAX Kynar Lined - 600 gpm H2O - 5 gpm acid SS304 1 $6,000 2009 strm.a200.214 136260 kg/hr 0.50 1.0 38801 0.28 $3,202 $3,223 $3,223

A- 203 Blowdown Tank Agitator Andritz Side-mounted, 3 x 75 hp. ( 170 kW) 170 kW 316LSS 1 INCLUDED

A- 204 Flash Tank Agitator UET Mixers Side-mounted, 3 x 75 hp. ( 170 kW) 170 kW 316LSS 3 $90,000 2009 strm.a200.254 252891 kg/hr 0.50 1.5 278194 1.10 $94,395 $95,028 $142,542

A- 208 Oligomer Hold Tank Agitator UET Mixers Side-mounted, 3 x 75 hp. ( 170 kW) 170 kW 316LSS 3 $90,000 2009 strm.a200.222 264116 kg/hr 0.50 1.5 292060 1.11 $94,641 $95,276 $142,914

A- 209 Ammonia Addition Tank Agitator Lotus 10 hp SS 1 $21,900 2009 strm.a200.228 410369 kg/hr 0.50 1.5 429554 1.05 $22,406 $22,556 $33,835

A- 210 Ammonia Static Mixer KOMAX SS 1 $5,000 2009 strm.a200.275 157478 kg/hr 0.50 1.0 151360 0.96 $4,902 $4,935 $4,935

A- 224 Reacidification Tank Agitator Lotus 20 hp SS 1 $0 2009 strm.a200.239 410846 kg/hr 0.50 1.5 429554 1.05 $0 $0 $0

C- 201 Transfer Conveyor Andritz 800 mm x 7600 mm (2'-8" x 25') 40 kW ea. AL-6XN 2 INCLUDED

C- 202 Distribution Conveyor Andritz 800 mm x 7600 mm (2'-8" x 25') 40 kW ea. AL-6XN 2 INCLUDED

C- 203 Overfeed Conveyor Andritz 800 mm x 7600 mm (2'-8" x 25') 40 kW ea. AL-6XN 4 INCLUDED

C- 204 Pressurized Heating Screw Andritz 2500 mm x 9500 mm (8' x 31') 75 kW Dup. 2205 1 INCLUDED

C- 205 Pressurized Pre-heater Discharge Andritz 850 mm x 3500 mm (2'-10" x 12') 40 kW ea. Dup. 2205 2 INCLUDED

C- 206 Pressurized Transport - No. 1 Andritz 900 mm x 3500 mm (3' x 12') 20 kW Incoloy 825-

CLAD; CS 1 INCLUDED

C- 207 Pressurized Transport - No. 2 Andritz 1200 mm x 3500 mm (4' x 12') 40 kW Incoloy 825-

CLAD; CS 1 INCLUDED

H- 201 Pretreatment Water Heater Mueller 29.9 MMBtu 304SS 1 $92,000 2010 Heat.A200.QH201 -8 Gcal/hr 0.70 2.2 -9 1.15 $101,545 $97,091 $213,600

H- 244 Waste Vapor Condenser Mueller Copied H-201 304SS 1 $34,000 2009 Heat.A200.QH244 2 Gcal/hr 0.70 2.2 7 3.98 $89,346 $89,946 $197,880

M-201 Doffing Roll Storage Bins Andritz 60 cu. m. (2100 cu.ft.) with conveyors/scrapers 40 kW ea. 304/316SS 2 INCLUDED

M-202 Pin Drum Feeder Andritz 15 kW ea. 316LSS 2 INCLUDED

M-203 Plug Screw Feeder Andritz 1100kW 

ea. 316LSS 2 INCLUDED

M-204 Prehydrolysis / Vertical Preheater Andritz 16' x 62' - 10 min. residence time 45 kW Dup. 2205 1 INCLUDED

M-205 Pin Drum Feeder Andritz 15 kW ea. Incoloy 825-

CLAD; CS 2 INCLUDED

M-206 Plug Screw Feeder Andritz 1100 kW 

ea.

Incoloy 825-

CLAD; CS 2 INCLUDED

M-207 Pretreatment Reactor Andritz 2600 mm x 9000 mm (9' x 30') - 2 min. residence 

time

Incoloy 825-

CLAD; CS 3 $19,812,400 2009 DRY101 83333 kg/hr 0.60 1.5 83333 1.00 $19,812,448 $19,945,315 $29,917,973

P- 201 Sulfuric Acid Pump Goulds 9 GPM, 245 FT TDH 316SS 1 $8,000 2009 strm.710 3720 kg/hr 0.80 2.3 1981 0.53 $4,832 $4,864 $11,187

P- 203 Blowdown Tank Discharge Pump Viking 1900 GPM, 150 FT TDH 125 316SS 1 $25,635 2010 strm.a200.222 292407 kg/hr 0.80 2.3 292060 1.00 $25,611 $24,487 $56,321

P- 204 Flash Tank Discharge Pump Viking 900 GPM, 150 FT TDH 75 316SS 1 $30,000 2009 strm.a200.254 204390 kg/hr 0.80 2.3 278194 1.36 $38,391 $38,649 $88,892

P- 208 Oligomer Hold Tank Discharge Viking 900 GPM, 150 FT TDH 75 316SS 1 $17,408 2010 strm.a200.223 292407 kg/hr 0.80 2.3 292060 1.00 $17,391 $16,629 $38,246

P- 209 Hydrolyzate Pump Goulds 1771 GPM, 150 FT TDH 100 316SS 1 $22,500 2009 strm.a200.228 402194 kg/hr 0.80 2.3 429554 1.07 $23,716 $23,875 $54,913

P- 239 Reacidified Hydrolyzate Pump Goulds 1771 GPM, 100 FT TDH 60 316SS 1 $0 2009 strm.a200.239 402194 kg/hr 0.80 2.3 429554 1.07 $0 $0 $0

S- 205 Hydrolyzate Solid-Liquid Separator Larox $35,000,000 2009 scis.a200.254b 39000 kg/hr 0.70 1.7 0 0.00 $0 $0 $0

T- 201 Sulfuric Acid Tank 12,800 gal, 24hr residence time PLASTIC 1 $6,210 2010 strm.710 1981 kg/hr 0.70 3.0 1981 1.00 $6,209 $5,937 $17,810

T- 203 Blowdown Tank Andritz 23' x 48' - 25 min. - 110,000 gal. SS316 1 INCLUDED

T- 204 Flash Tank 23' x 48' - 110,000 gal. SS316 1 $511,000 2009 strm.a200.223 264116 kg/hr 0.70 2.0 292060 1.11 $548,271 $551,948 $1,103,895

T- 208 Oligomer Conversion Tank 2.6 atm, 130C operating

30 min. hold = 30,000 gal SS316 1 $203,000 2009 strm.a200.223 264116 kg/hr 0.70 2.0 292060 1.11 $217,806 $219,267 $438,534

T- 209 Ammonia Addition Tank 118,000 gal, 1hr residence time SS304 1 $236,000 2009 strm.a200.228 410369 kg/hr 0.70 2.0 429554 1.05 $243,670 $245,304 $490,609

T- 224 Reacidification Tank Mueller 500,000 gallon SS304 1 $0 2009 strm.a200.239 410369 kg/hr 0.70 2.0 429554 1.05 $0 $0 $0

Area 200 Totals $21,348,782 $21,484,330 $32,957,310

EQPT NO
99 

Mechanical Equipment List Scaled Installed Costs

EQUIPMENT TITLE VENDOR DESCRIPTION HP MATERIAL

NUM REQD

$

Year of 

Quote

Scaling Variable

Scaling Val

Units

Scaling Exp

Inst Factor

New Val

Size Ratio

Scaled Purch 

Cost

Purch Cost in 

Proj year

Inst Cost in Proj 

year

A- 300 Ethanol Fermentor Agitator Lotus 30 hp SS304 1 $52,500 2009 inumsscf 1 ea 1.00 1.5 12 12 $630,000 $634,225 $951,337

A- 301 Seed Hold Tank Agitator Lotus 15 hp SS304 1 $31,800 2009 STRM.A300.304 40414 kg/hr 0.50 1.5 42607 1.05 $32,651 $32,870 $49,305

A- 304 4th Seed Vessel Agitator Lotus 7.5 hp SS 2 $26,000 2009 2 ea 0.50 1.5 2 1.00 $26,000 $26,174 $39,262

A- 305 5th Seed Vessel Agitator Lotus 10 hp SS 2 $43,000 2009 2 ea 0.50 1.5 2 1.00 $43,000 $43,288 $64,933

A- 306 Beer Surge Tank Agitator Lotus 20 hp SS304 2 $68,300 2009 strm.501 425878 kg/hr 0.50 1.5 450740 1.06 $70,265 $70,737 $106,105

A- 308 Enzyme-Hydrolysate Mixer GLV inline mixer 1673 gpm 100 hp SS316 1 $109,000 2009 strm.a300.t310fd 379938 kg/hr 0.50 1.7 443391 1.17 $117,751 $118,540 $201,519

F- 300 Ethanol Fermentor Mueller 1,000,000 gallon ea 304SS 12 $10,128,000 2009 inumsscf 12 ea 1.00 1.5 12 1.00 $10,128,000 $10,195,921 $15,293,882

F- 301 1st Seed Fermentor A&B Process 20 gallon skid complete - $37,700 ea 304SS 2 $75,400 2009 2 ea 0.70 1.8 2 1.00 $75,400 $75,906 $136,630

F- 302 2nd Seed Fermentor A&B Process 200 gallon skid complete - $58,300 ea 304SS 2 $116,600 2009 2 ea 0.70 1.8 2 1.00 $116,600 $117,382 $211,288

F- 303 3rd Seed Fermentor A&B Process 2000 gallon skid complete - $78,800 ea 304SS 2 $157,600 2009 2 ea 0.70 1.8 2 1.00 $157,600 $158,657 $285,582

F- 304 4th Seed Fermentor Mueller 20,000 gallon, incl. coil - $176,000 ea 304SS 2 $352,000 2009 2 ea 0.70 2.0 2 1.00 $352,000 $354,361 $708,721

F- 305 5th Seed Fermentor Mueller 200,000 gallon, incl. coil - $590,000 ea 304SS 2 $1,180,000 2009 2 ea 0.70 2.0 2 1.00 $1,180,000 $1,187,913 $2,375,827

H- 300 Fermentation Cooler Alfa Laval Plate & frame 304SS 12 $86,928 2009 inumsscf 12 ea 1.00 2.2 12 1.00 $86,928 $87,511 $192,524

H- 301 Hydrolyzate Cooler Fox Engineering Plate & Frame 32.5 MMBtu/hr SS 304 1 $85,000 2010 heat.a300.qc301 8 Gcal/hr 0.70 2.2 11 1.30 $101,972 $97,500 $214,500

H- 304 4th Seed Fermentor Coil Mueller incl. w/ tank 304SS 1 INCLUDED

H- 305 5th Seed Fermentor Coil Mueller incl. w/ tank 304SS 1 INCLUDED

H- 310 Fermentor Batch Cooler Alfa Laval Plate & frame SS304 1 $23,900 2009 heat.a300.qf300sf 5 Gcal/hr 0.70 1.8 0.000 0.00 $0 $0 $0

P- 300 Fermentation Recirc/Transfer Pump Goulds 340 GPM, 150 FT 20 316SS 5 $47,200 2009 inumsscf 12 ea 0.80 2.3 12 1.00 $47,200 $47,517 $109,288

P- 301 Seed Hold Transfer Pump Goulds 190 GPM, 150 FT TDH 10 316SS 1 $8,200 2009 strm.a300.304 43149 kg/hr 0.80 2.3 42607 0.99 $8,117 $8,172 $18,795

P- 302 Seed Transfer Pump Goulds 190 GPM, 615 FT TDH 40 316SS 2 $24,300 2009 strm.a300.304 43149 kg/hr 0.80 2.3 42607 0.99 $24,055 $24,217 $55,698

P- 306 Beer Transfer Pump Goulds 2152 GPM, 171 FT TDH 125 316SS 1 $26,800 2009 strm.501 488719 kg/hr 0.80 2.3 450740 0.92 $25,120 $25,289 $58,165

P- 310 Saccharification Transfer Pump Goulds 352 GPM, 150 FT TDH 20 316SS 5 $47,200 2009 strm.a300.306 421776 kg/hr 0.80 2.3 443391 1.05 $49,125 $49,455 $113,746

T- 301 Seed Hold Tank Mueller 300,000 gallon 316SS 1 $439,000 2009 strm.a300.304 40414 kg/hr 0.70 1.8 42607 1.05 $455,540 $458,595 $825,471

T- 306 Beer Storage Tank Mueller 500,000 gallon 316SS 1 $636,000 2009 strm.501 425878 kg/hr 0.70 1.8 450740 1.06 $661,768 $666,206 $1,199,170

T- 310 Saccharification Tank Caldwell 250,000 gal each - 19' dia. x 120' tall 304SS 8 $3,840,000 2009 strm.a300.306 421776 kg/hr 0.70 2.0 443391 1.05 $3,976,717 $4,003,386 $8,006,772

Area 300 Totals $18,365,811 $18,483,821 $31,218,520

A- 400 Cellulase Fermentor Agitators Lotus 800.0 SS316 $580,000 2009 CLVESSEL 1 ea 1.00 1.5 9 9.00 $5,220,000 $5,255,007 $7,882,510

A- 401 Cellulase Fermentor Agitators Lotus 0.75hp SS316 $3,420 2009 ICLSEED 1 ea 1.00 1.5 4 4.00 $13,680 $13,772 $20,658

A- 402 Cellulase Fermentor Agitators Lotus 8 hp SS316 $63,000 2009 ICLSEED 1 ea 1.00 1.5 4 4.00 $252,000 $253,690 $380,535

A- 403 Cellulase Fermentor Agitators Lotus 80 hp SS316 $11,000 2009 ICLSEED 1 ea 1.00 1.5 4 4.00 $44,000 $44,295 $66,443

A- 405 Media-Prep Tank Agitator Lotus 7.5 hp A285C 1 $8,500 2009 strm.a400.402a 12255 kg/hr 0.50 1.5 0 0.00 $0 $0 $0

A- 406 Cellulase Nutrient Mix Tank Agitator Lotus 3 hp CS 1 $4,800 2009 strm.a400.416 174 kg/hr 0.50 1.6 224 1.29 $5,446 $5,482 $8,772

A- 410 Cellulase Hold Tank Agitator Lotus 10 hp SS316 1 $26,900 2009 strm.422 10930 kg/hr 0.50 1.5 13836 1.27 $30,266 $30,469 $45,703

F- 400 Cellulase Fermentor Mueller 80,000 gal, 1 atm, 28 °C, Internal coil SS316 $400,500 2009 CLVESSEL 1 ea 1.00 2.0 9 9.00 $3,604,500 $3,628,673 $7,257,345

F- 401 1st Cellulase Seed Fermentor A&B Process 80 gallon skid complete - $46,000 ea 304SS $46,000 2009 ICLSEED 1 ea 1.00 1.8 4 4.00 $184,000 $185,234 $333,421

F- 402 2nd Cellulase Seed Fermentor A&B Process 800 gallon skid complete - $57,500 ea 304SS $57,500 2009 ICLSEED 1 ea 1.00 1.8 4 4.00 $230,000 $231,542 $416,776

F- 403 3rd Cellulase Seed Fermentor A&B Process 8,000 gallon skid complete - $95,400 ea 304SS $95,400 2009 ICLSEED 1 ea 1.00 1.8 4 4.00 $381,600 $384,159 $691,486

H- 400 Cellulase Fermentation Cooler Mueller Cooling coil included with Cellulase Fermenter 304SS INCLUDED

H- 405 Media Prep Tank Cooler Mueller Cooling coil included with Media Prep Tank 304SS 1 INCLUDED

M-401 Fermenter Air Compressor Package Dresser Roots 8000 SCFM @ 16 psig CS 2 $350,000 2009 strm.a400.450 33168 kg/hr 0.60 1.6 32583 0.98 $346,282 $348,604 $557,767

P- 400 Cellulase Transfer Pump Goulds 59 gpm, 100 FT, TDH SIZE 2X1-10C 3 316SS 1 $7,357 2010 strm.a400.420 13399 kg/hr 0.80 2.3 13836 1.03 $7,549 $7,218 $16,600

P- 401 Cellulase Seed Pump Goulds 3 GPM, 100 FT TDH SIZE 2X1-10 2 316SS 4 $29,972 2010 strm.a400.409 681 kg/hr 0.80 2.3 770 1.13 $33,045 $31,596 $72,670

P- 405 Media Pump Goulds 63 GPM, 100 FT TDH SIZE 2X1-10C 3 316SS 1 $7,357 2010 strm.a400.402a 14307 kg/hr 0.80 2.3 0 0.00 $0 $0 $0

P- 406 Cellulase Nutrient Transfer Pump Viking Gear Pump 2 GPM, 100 FT 1 316SS 1 $1,500 2009 strm.a400.416 454 kg/hr 0.80 2.3 224 0.49 $852 $858 $1,974

P- 410 Cellulase Feed Pump Viking Gear Pump 1 316SS 1 $1,500 2009 strm.a400.422 18168 kg/hr 0.80 2.3 13836 0.76 $4,584 $4,615 $10,614

P- 420 Anti-foam Pump Viking Gear Pump 2 GPM, 100 FT 1 316SS 1 $1,500 2009 strm.a400.444 11 kg/hr 0.80 2.3 13.4 1.26 $1,805 $1,817 $4,179

T- 401 SO2 Storage Tank SO2 Vendor 1 ton cylinders, incl w/ delivery SS304 1 $0

T- 405 Media-Prep Tank Mueller 20,000 gallon, incl. coil 304SS 1 $176,000 2009 strm.a400.402a 12255 kg/hr 0.70 1.8 0 0.00 $0 $0 $0

T- 406 Cellulase Nutrient Mix Tank Harrington HDPE, 8,000 gal HDPE 1 $9,000 2010 strm.a400.416 224 kg/hr 0.70 3.0 224 1.00 $9,000 $8,605 $25,815

T- 410 Cellulase Hold Tank Mueller 80,000 gal 304SS 1 $248,070 2009 strm.a400.422 10930 kg/hr 0.70 1.8 13836 1.27 $292,589 $294,551 $530,192

Area 400 Totals $10,661,197 $10,730,186 $18,323,460

EQPT NO
100 

Mechanical Equipment List Scaled Installed Costs

EQUIPMENT TITLE VENDOR DESCRIPTION HP MATERIAL

NUM REQD

$

Year of 

Quote

Scaling Variable

Scaling Val

Units

Scaling Exp

Inst Factor

New Val

Size Ratio

Scaled Purch 

Cost

Purch Cost in 

Proj year

Inst Cost in Proj 

year

A- 530 Filtrate Tank Agitator Lotus 7.5 hp SS 1 $26,000 2009 strm.a500.evap.572 337439 kg/hr 0.50 1.5 355024 1.05 $26,669 $26,848 $40,272

C- 501 Lignin Wet Cake Conveyor KWS/Barnard-Boe Belt 100 ft. long x 24" wide, enclosed 10 SS304 1 $70,000 2009 strm.a500.571 28630 kg/hr 0.80 1.7 36538 1.28 $85,082 $85,653 $145,610

C- 502 Lignin Wet Cake Screw KWS/Barnard-Boe Screw conveyor - 25 ft lg x 14" dia 15 SS304 1 $20,000 2009 strm.a500.571 28630 kg/hr 0.80 1.7 36538 1.28 $24,309 $24,472 $41,603

D- 501 Beer Column Megtec 14' dia. x 76' tall, 32 trays, 24" spacing 316SS 1 $3,407,000 2009 strm.a500.511 30379 kg/hr 0.60 2.4 29213 0.96 $3,327,914 $3,350,232 $8,040,557

D- 502 Rectification Column Megtec Top 13' dia. x 68' h, Bottom 4'6" x 31' h 316SS 1 INCLUDED

H- 501 Beer Column Reboiler Megtec S & T 316SS;CS 1 INCLUDED

H- 502 Rectification Column Reboiler Megtec S & T 316SS;CS 1 INCLUDED

H- 504 Beer Column Condenser Megtec S & T 316SS;CS 1 INCLUDED

H- 505 Rectification Column Condenser GEA Rainey 92.2 MM Btu/hr 3 cells 300 tot CS 1 $487,000 2010 heat.a500.qcd502 23 Gcal/hr 0.60 2.8 23 1.00 $486,746 $465,399 $1,303,116

H- 512 Beer Column Feed Interchanger Megtec Plate & Frame 316SS 1 INCLUDED

M-503 Molecular Sieve Package (9 pieces) Delta-T SS 1 $2,601,000 2009 strm.a500.515 22687 kg/hr 0.60 1.8 21808 0.96 $2,540,057 $2,557,091 $4,602,764

M-505 Pressure Filter Pressing Compr Atlas-Copco 460 SCFM, 300 psig 150 hp 1 $75,200 2009 strm.a500.evap.sqairin 808 kg/hr 0.60 1.6 809 1.00 $75,245 $75,750 $121,199

M-507 Pressure Filter Drying Compr Atlas-Copco 4000 SCFM, 130 psig (ea) 700 hp ea. 2 $405,000 2009 strm.a500.evap.557 12233 kg/hr 0.60 1.6 12105 0.99 $402,453 $405,152 $648,243

P- 501 Beer Column Bottoms Recirc Pump Megtec 10,000 gpm 200 316SS 1 INCLUDED

P- 503 Beer Column Reflux Pump Megtec 15 gpm 2 316SS 1 INCLUDED

P- 504 Rectification Column Btms Pump Megtec 150 gpm 15 316SS 1 INCLUDED

P- 505 Rectification Column Reflux Pump Megtec 500 gpm 50 316SS 1 INCLUDED

P- 506 Beer Column Stillage Pump Megtec 2000 gpm 75 316SS 1 INCLUDED

P- 515 Scrubber Bottoms Pump Goulds 108 GPM, 104 FT TDH 316SS 1 $6,300 2009 strm.551 24527 kg/hr 0.80 2.3 27197 1.11 $6,843 $6,889 $15,844

P- 530 Filtrate Tank Discharge Pump Sulzer 590 GPM, 100 FT TDH SIZE 4X3-13 SS 1 $13,040 2010 strm.a500.evap.571 31815 kg/hr 0.80 2.3 36538 1.15 $14,567 $13,928 $32,035

P- 531 Feed Pump Warman 1014 GPM 230 FT TDH SIZE 8X6-15 100 hp SS 1 $18,173 2010 strm.a500.evap.571 31815 kg/hr 0.80 2.3 36538 1.15 $20,301 $19,411 $44,645

P- 532 Manifold Flush Pump Warman 100 hp SS 1 $17,057 2010 strm.a500.evap.571 31815 kg/hr 0.80 2.3 36538 1.15 $19,054 $18,219 $41,903

P- 533 Cloth Wash Pump Warman 150 hp SS 1 $29,154 2010 strm.a500.evap.571 31815 kg/hr 0.80 2.3 36538 1.15 $32,568 $31,140 $71,621

P- 581 Filtrate Discharge Pump Sulzer 590 GPM, 100 FT TDH SIZE 4X3-13 75 hp SS 1 $13,040 2010 strm.a500.evap.571 31815 kg/hr 0.80 2.3 36538 1.15 $14,567 $13,928 $32,035

S- 505 Pressure Filter Larox 384 sq. m fitration area ea incl packing SS316 2 $3,294,700 2010 strm.a500.evap.571 31815 kg/hr 0.80 1.7 36538 1.15 $3,680,519 $3,519,098 $5,982,467

T- 503 Beer Column Reflux Drum Megtec 4' dia, 6' high, 50 psig design 316SS 1 INCLUDED

T- 505 Rectification Column Reflux Drum Megtec 4' dia, 6' high, 50 psig design 316SS 1 INCLUDED

T- 512 Vent Scrubber Envitech Inlet Gas: 9681 acfm, 91°F, 1.97 mass% Ethanol SS304;PP 1 $215,000 2009 strm.a500.523 22608 kg/hr 0.60 2.4 21759 0.96 $210,121 $211,530 $507,673

T- 530 Filtrate Tank 13,750 gal 14' dia x 12' H SS 1 $103,000 2010 strm.a500.evap.571 31815 kg/hr 0.70 2.0 36538 1.15 $113,480 $108,503 $217,006

T- 531 Feed Tank 20,300 gal 14' dia x 18' H SS 1 $174,800 2010 strm.a500.evap.571 31815 kg/hr 0.70 2.0 36538 1.15 $192,585 $184,139 $368,278

T- 532 Recycled Water Tank Harrington Plastic 4000 gal. HDPE 1 $1,520 2010 strm.a500.evap.571 31815 kg/hr 0.70 3.0 36538 1.15 $1,675 $1,601 $4,804

T- 533 Pressing Air Compressor Receiver 1350 gal., 300 psig design CS 1 $8,000 2010 strm.a500.evap.571 31815 kg/hr 0.70 3.1 36538 1.15 $8,814 $8,427 $26,125

T- 534 Drying Air Compressor Receiver 9,000 gal., 150 psig design CS 2 $17,000 2010 strm.a500.evap.571 31815 kg/hr 0.70 3.1 36538 1.15 $18,730 $17,908 $55,516

Area 500 Totals $11,302,300 $11,145,318 $22,343,314

EQPT NO
101 

Mechanical Equipment List Scaled Installed Costs

EQUIPMENT TITLE VENDOR DESCRIPTION HP MATERIAL

NUM REQD

$

Year of 

Quote

Scaling Variable

Scaling Val

Units

Scaling Exp

Inst Factor

New Val

Size Ratio

Scaled Purch 

Cost

Purch Cost in 

Proj year

Inst Cost in Proj 

year

B- 606 Biogas Blower ADI 3000 ACFM, 10 PSI 200.0 4 INCLUDED

B- 608 Aerobic Digester Blower ADI 14,000 SCFM @ 10.3 psig 1,000 8 $1,933,750 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $1,986,627 $1,899,498 $1,899,498

C- 614 Aerobic Sludge Screw CS 1 $25,000 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $25,684 $24,557 $24,557

H- 602 Anaerobic Digestor Feed Cooler ALFA LAVAL Wide gap plate & frame 2.5 mgd SS316; CS 1 $83,863 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $86,156 $82,378 $82,378

M-606 Biogas Emergency Flare ADI 4 $32,955 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $33,856 $32,371 $32,371

M-630 Polymer Addition System ADI 11.4 gph neat polymer 10.0 1 $9,300 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $9,554 $9,135 $9,135

M-632 Caustic Feed System ADI 0-300 gph Aerobic Digesters #1- #3 1.5 3 $22,800 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $23,423 $22,396 $22,396

M-640 Evaporator System $3,801,095 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $3,905,034 $3,733,767 $3,733,767

P- 602 Anaerobic Reactor Feed Pump ADI 2500 gpm submersible rail mounted 50.0 CS 4 $231,488 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $237,818 $227,388 $227,388

P- 606 Anaerobic Reactor Recirc Pump ADI 400 gpm, 50 ft TDH 7.5 4 INCLUDED

P- 607 Waste Anaerobic Sludge Pump ADI 10 gpm 3.0 6 $93,300 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $95,851 $91,647 $91,647

P- 608 Aeration Basin Feed Pump ADI 15.0 4 $84,000 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $86,297 $82,512 $82,512

P- 609 Return Activated Sludge Pump ADI 40.0 6 $177,300 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $182,148 $174,160 $174,160

P- 610 Reverse Osmosis Feed Pump ADI ? 1 INCLUDED

P- 611 Centrifuge Feed Pump ADI 10.0 2 $61,200 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $62,873 $60,116 $60,116

P- 612 Centrate Pump ADI 15.0 2 $70,800 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $72,736 $69,546 $69,546

P- 616 Treated Water Pump CS 1 INCLUDED

R- 609 Membrane Bioreactor ADI Includes membrane CIP and Scour system 3 $5,248,750 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $5,392,274 $5,155,780 $5,155,780

S- 610 Reverse Osmosis System ADI 1 $2,210,979 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $2,271,437 $2,171,816 $2,171,816

S- 611 Centrifuge 3 $6,493,500 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $6,671,062 $6,378,482 $6,378,482

T- 606 Anaerobic Basin ADI 31 Million gallonwith cover CONCRETE 4 $27,000,000 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $27,738,302 $26,521,754 $26,521,754

T- 608 Aeration Digester ADI 25 ft H x 115 ft x 344 ft 2 ft floor/wall thkness CONCRETE 3 $2,700,000 2010 strm.a600.601 393100 kg/hr 0.60 1.0 411178 1.05 $2,773,830 $2,652,175 $2,652,175

T- 609 Sludge Holding Tank ADI 1 INCLUDED

Area 600 Totals $51,654,964 $49,389,478 $49,389,478

A- 701 Denaturant In-line Mixer 4 inch SS304 1 $3,850 2009 strm.703 23154 kg/hr 0.50 1.0 22273 0.96 $3,776 $3,801 $3,801

A- 720 CSL Storage Tank Agitator Lotus 10 hp SS304 1 $21,200 2009 strm.735 1393 kg/hr 0.50 1.5 1323 0.95 $20,660 $20,798 $31,198

A- 760 DAP Make-up Tank Agitator Lotus 5.5 hp SS304 1 $9,800 2009 strm.755 163 kg/hr 0.50 1.5 142 0.87 $9,152 $9,214 $13,820

C- 755 DAP Bulk Bag Unloader Flexicon Super sack unloader 1 $30,000 2009 strm.755 163 kg/hr 0.60 1.7 142 0.87 $27,636 $27,822 $47,297

P- 701 Ethanol Product Pump Goulds 150 GPM, 112 FT TDH 5.0 CS 2 $9,200 2009 strm.a500.515 22681 kg/hr 0.80 3.1 21808 0.96 $8,916 $8,975 $27,824

P- 703 Sulfuric Acid Pump Goulds 5 GPM, 150 FT TDH SIZE 2X1-10 0.5 SS316 1 $7,493 2010 strm.710 1981 kg/hr 0.80 2.3 1981 1.00 $7,492 $7,163 $16,475

P- 704 Firewater Pump Goulds 2500 GPM, 150 FT TDH 125.0 CS 1 $15,000 2009 strm.a700.713 8343 kg/hr 0.80 3.1 8021 0.96 $14,536 $14,633 $45,362

P- 710 Gasoline Pump Goulds 4 GPM, 60 FT 0.5 CS 1 $3,000 2009 strm.a700.701 473 kg/hr 0.80 3.1 465 0.98 $2,959 $2,979 $9,234

P- 720 CSL Pump Goulds 8 GPM, 80 FT TDH 0.5 CS 1 $3,000 2009 strm.735 1393 kg/hr 0.80 3.1 1323 0.95 $2,879 $2,898 $8,984

P- 760 DAP Pump Goulds 2 GPM, 100 FT TDH 0.5 CS 1 $3,000 2009 strm.755 163 kg/hr 0.80 3.1 142 0.87 $2,689 $2,707 $8,392

T- 701 Ethanol Product Storage Tank Mueller 750,000 gal., 7 day storage, Floating roof A285C 2 $1,340,000 2009 strm.a500.515 22681 kg/hr 0.70 1.7 21808 0.96 $1,303,683 $1,312,426 $2,231,125

T- 703 Sulfuric Acid Storage Tank 12,600 gal, 12' dia x15' H SS 1 $96,000 2010 strm.710 1981 kg/hr 0.70 1.5 1981 1.00 $95,987 $91,777 $137,666

T- 704 Firewater Storage Tank 600,000 gal - 4 hrs @ 2500 gpm Glass lined 1 $803,000 2009 strm.a700.713 8343 kg/hr 0.70 1.7 8021 0.96 $781,201 $786,440 $1,336,948

T- 706 Ammonia Storage Tank Chemithon 28,000 gal SA- 516-70 2 $196,000 2010 strm.717 1171 kg/hr 0.70 2.0 1166 1.00 $195,389 $186,820 $373,639

T- 710 Gasoline Storage Tank 65,000 gal., floating roof CS 1 $200,000 2009 strm.a700.701 473 kg/hr 0.70 1.7 465 0.98 $197,602 $198,927 $338,176

T- 720 CSL Storage Tank 70,000 gal Glass lined CS 1 $70,000 2009 strm.735 1393 kg/hr 0.70 2.6 1323 0.95 $67,516 $67,969 $176,720

T- 755 DAP Bulk Bag Holder Flexicon Super sack holder 1 INCLUDED

T- 760 DAP Make-up Tank 12,800 gal SS304 1 $102,000 2009 strm.717 1615 kg/hr 0.70 1.8 1166 0.72 $81,192 $81,737 $147,126

Area 700 Totals $2,823,265 $2,827,086 $4,953,786

EQPT NO
102 

Mechanical Equipment List Scaled Installed Costs

EQUIPMENT TITLE VENDOR DESCRIPTION HP MATERIAL

NUM REQD

$

Year of 

Quote

Scaling Variable

Scaling Val

Units

Scaling Exp

Inst Factor

New Val

Size Ratio

Scaled Purch 

Cost

Purch Cost in 

Proj year

Inst Cost in Proj 

year

H- 801 Burner Combustion Air Preheater Babcock & Wilcox 1 INCLUDED

H- 811 BFW Preheater Babcock & Wilcox 1 INCLUDED

H- 812 Pretreatment/BFW heat recovery Mueller 9.4 MM Btu/hr SS304 1 $41,000 2009 heat.QH812 -2 Gcal/hr 0.70 2.2 -2 0.93 $38,856 $39,117 $86,057

M-802 Air Intake Fan Babcock & Wilcox INCLUDED

M-803 Boiler Babcock & Wilcox 525,000 lb/hr @ 900 psig 2752 kW CS 1 $28,550,000 2010 strm.a800.a810.812 238686 kg/hr 0.60 1.8 234784 0.98 $28,269,041 $27,029,216 $48,652,589

M-804 Combustion Gas Baghouse Babcock & Wilcox INCLUDED WITH FGD SYSTEM INCLUDED

M-811 Turbine/Generator Siemens 23.6 kW, 2 extractions 1 $9,500,000 2010 work.wtotal -42200 kW 0.60 1.8 -41324 0.98 $9,381,215 $8,969,773 $16,145,591

M-820 Hot Process Water Softener System Proctor Sales 1 $78,000 2010 strm.a800.a810.812 235803 kg/hr 0.60 1.8 234784 1.00 $77,798 $74,386 $133,894

M-830 Amine Addition Pkg. Proctor Sales 1 $40,000 2010 strm.a800.a810.812 235803 kg/hr 0.00 1.8 234784 1.00 $40,000 $38,246 $68,842

M-832 Ammonia Addition Pkg Proctor Sales 1 INCLUDED

M-834 Phosphate Addition Pkg. Proctor Sales 1 INCLUDED

P- 804 Condensate Pump Proctor Sales SS316 2 INCLUDED

P- 811 Turbine Condensate Pump Proctor Sales SS304 2 INCLUDED

P- 824 Deaerator Feed Pump Proctor Sales SS304 2 INCLUDED

P- 826 BFW Pump Proctor Sales SS316 5 INCLUDED

P- 828 Blowdown Pump Proctor Sales CS 2 INCLUDED

P- 830 Amine Transfer Pump Proctor Sales CS 1 INCLUDED

T- 804 Condensate Collection Tank Proctor Sales A285C 1 INCLUDED

T- 824 Condensate Surge Drum Proctor Sales SS304 1 INCLUDED

T- 826 Deaerator Proctor Sales Tray type CS;SS316 1 $305,000 2010 strm.a800.a810.812 235803 kg/hr 0.60 3.0 234784 1.00 $304,209 $290,867 $872,600

T- 828 Blowdown Flash Drum Proctor Sales CS 1 INCLUDED

T- 830 Amine Drum Proctor Sales SS316 1 INCLUDED

Area 800 Totals $38,111,118 $36,441,604 $65,959,573

M-902 Cooling Tower System Marley SPX 44,200 gpm 750 hp FIBERGLASS 1 $1,375,000 2010 strm.a900.945 10037820 kg/hr 0.60 1.5 11923904 1.19 $1,524,650 $1,457,782 $2,186,673

M-904 Plant Air Compressor Rogers Machinery 400 SCFM@125 psig 150 hp 1 $28,000 2010 DRY101 83333 kg/hr 0.60 1.6 83333 1.00 $28,000 $26,772 $42,835

M-908 Chilled Water Package Trane 2 x 2350 tons (14.2 MM kcal/hr) 3400 hp 1 $1,275,750 2010 heat.a900.qchwop 14 Gcal/hr 0.60 1.6 13 0.95 $1,234,354 $1,180,217 $1,888,348

M-910 CIP System 100,000 GAL SS304/SS316 1 $421,000 2009 strm.a900.914 63 kg/hr 0.60 1.8 145 2.30 $694,222 $698,878 $1,257,980

P- 902 Cooling Water Pump Goulds 16,120 GPM, 100 FT TDH SIZE 20X20-28 500.0 CS 3 $283,671 2010 strm.a900.945 10982556 kg/hr 0.80 3.1 11923904 1.09 $302,961 $289,674 $897,989

P- 912 Make-up Water Pump Goulds 685 GPM, 75 FT TDH SIZE 6X4-13 20.0 CS 1 $6,864 2010 strm.a900.904 155564 kg/hr 0.80 3.1 147140 0.95 $6,565 $6,277 $19,459

P- 914 Process Water Circulating Pump Goulds 2285 GPM, 75 FT TDH SIZE 8X6-13 75.0 CS 1 $15,292 2010 strm.a900.905 518924 kg/hr 0.80 3.1 523463 1.01 $15,399 $14,724 $45,643

S- 904 Instrument Air Dryer Zeks 670 SCFM - CYCLING TYPE CS 1 $15,000 2009 DRY101 83333 kg/hr 0.60 1.8 83333 1.00 $15,000 $15,101 $27,181

T- 904 Plant Air Receiver Rogers Machinery 3800 gal - 72" x 228" vertical CS 1 $16,000 2009 DRY101 83333 kg/hr 0.60 3.1 83333 1.00 $16,000 $16,107 $49,933

T- 914 Process Water Tank No. 1 250,000 gal CS 1 $250,000 2009 strm.a900.905 451555 kg/hr 0.70 1.7 523463 1.16 $277,245 $279,104 $474,476

Area 900 Totals $4,114,396 $3,984,635 $6,890,517

Grand Totals $158,381,833 $154,486,458 $232,035,959

EQPT NO
103 

Appendix B. Discounted Cash Flow Rate of Return Worksheet
104 

Year -2 -1 0 1 2 3 4 5 6

Fixed Capital Investment $12,819,631 $96,147,231 $51,278,523

Land $1,848,000

Working Capital $20,030,673

Loan Payment $35,821,932 $35,821,932 $35,821,932 $35,821,932 $35,821,932 $35,821,932

 Loan Interest Payment $1,538,356 $13,076,023 $19,229,446 $19,229,446 $17,902,047 $16,468,457 $14,920,179 $13,248,038 $11,442,127

 Loan Principal $19,229,446 $163,450,293 $240,368,078 $223,775,593 $205,855,708 $186,502,233 $165,600,480 $143,026,587 $118,646,782

 Ethanol Sales $114,997,260 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440

 By-Product Credit $5,745,024 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742

Total Annual Sales $120,742,284 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181

Annual Manufacturing Cost

 Feedstock $39,550,982 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122

 Baghouse Bags $466,183 $466,183

 Other Variable Costs $25,025,596 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969

 Fixed Operating Costs $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826

Total Product Cost $75,698,586 $82,550,916 $82,550,916 $82,550,916 $82,550,916 $83,017,099

Annual Depreciation

General Plant Writedown 14% 24.49% 17.49% 12.49% 8.93% 8.92%

 Depreciation Charge $47,822,041 $81,956,738 $58,530,965 $41,798,271 $29,884,592 $29,851,127

 Remaining Value $286,831,849 $204,875,112 $146,344,146 $104,545,875 $74,661,283 $44,810,156

Steam Plant Writedown 3.75% 7.22% 6.68% 6.18% 5.71% 5.29%

 Depreciation Charge $2,473,484 $4,761,622 $4,404,121 $4,074,323 $3,768,270 $3,485,963

 Remaining Value $63,486,089 $58,724,468 $54,320,347 $50,246,024 $46,477,754 $42,991,790

Net Revenue ($24,481,274) ($49,180,142) ($23,963,278) ($5,352,508) $8,539,364 $10,194,864

Losses Forward ($24,481,274) ($73,661,416) ($97,624,694) ($102,977,201) ($94,437,838)

Taxable Income ($24,481,274) ($73,661,416) ($97,624,694) ($102,977,201) ($94,437,838) ($84,242,973)

Income Tax $0 $0 $0 $0 $0 $0

Annual Cash Income $9,221,766 $19,618,333 $19,618,333 $19,618,333 $19,618,333 $19,152,150

Discount Factor 1.2100 1.1000 1.0000 0.9091 0.8264 0.7513 0.6830 0.6209 0.5645

Annual Present Value $227,041,341 $8,383,423 $16,213,498 $14,739,544 $13,399,585 $12,181,441 $10,810,889

Total Capital Investment + Interest $19,609,244 $120,145,580 $90,538,643

Net Present Worth $0

Year 7 8 9 10 11 12 13 14 15 16

Fixed Capital Investment

Land

Working Capital

Loan Payment $35,821,932 $35,821,932 $35,821,932 $35,821,932 $0 $0 $0 $0 $0 $0

 Loan Interest Payment $9,491,743 $7,385,327 $5,110,399 $2,653,476 $0 $0 $0 $0 $0 $0

 Loan Principal $92,316,592 $63,879,988 $33,168,455 $0 $0 $0 $0 $0 $0 $0

 Ethanol Sales $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440

 By-Product Credit $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742

Total Annual Sales $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181

Annual Manufacturing Cost

 Feedstock $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122

 Baghouse Bags $466,183 $466,183

 Other Variable Costs $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969

 Fixed Operating Costs $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826

Total Product Cost $82,550,916 $82,550,916 $82,550,916 $82,550,916 $83,017,099 $82,550,916 $82,550,916 $82,550,916 $82,550,916 $83,017,099

Annual Depreciation

General Plant Writedown 8.93% 4.46%

 Depreciation Charge $29,884,592 $14,925,564

 Remaining Value $14,925,564 $0

Steam Plant Writedown 4.89% 4.52% 4.46% 4.46% 4.46% 4.46% 4.46% 4.46% 4.46% 4.46%

 Depreciation Charge $3,224,104 $2,982,692 $2,943,116 $2,942,457 $2,943,116 $2,942,457 $2,943,116 $2,942,457 $2,943,116 $2,942,457

 Remaining Value $39,767,686 $36,784,995 $33,841,878 $30,899,422 $27,956,306 $25,013,849 $22,070,733 $19,128,276 $16,185,160 $13,242,704

Net Revenue $12,839,826 $30,146,682 $47,386,750 $49,844,332 $52,030,966 $52,497,808 $52,497,149 $52,497,808 $52,497,149 $52,031,625

Losses Forward ($84,242,973) ($71,403,147) ($41,256,465) $0 $0 $0 $0 $0 $0 $0

Taxable Income ($71,403,147) ($41,256,465) $6,130,284 $49,844,332 $52,030,966 $52,497,808 $52,497,149 $52,497,808 $52,497,149 $52,031,625

Income Tax $0 $0 $2,145,599 $17,445,516 $18,210,838 $18,374,233 $18,374,002 $18,374,233 $18,374,002 $18,211,069

Annual Cash Income $19,618,333 $19,618,333 $17,472,734 $2,172,817 $36,763,244 $37,066,032 $37,066,263 $37,066,032 $37,066,263 $36,763,013

Discount Factor 0.5132 0.4665 0.4241 0.3855 0.3505 0.3186 0.2897 0.2633 0.2394 0.2176

Annual Present Value $10,067,307 $9,152,097 $7,410,145 $837,715 $12,885,293 $11,810,380 $10,736,776 $9,760,645 $8,873,369 $8,000,703

Total Capital Investment + Interest

Net Present Worth
105 

Year 17 18 19 20 21 22 23 24 25 26

Fixed Capital Investment

Land

Working Capital

Loan Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0

 Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0

 Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0

 Ethanol Sales $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440 $131,425,440

 By-Product Credit $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742 $6,565,742

Total Annual Sales $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181 $137,991,181

Annual Manufacturing Cost

 Feedstock $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122 $45,201,122

 Baghouse Bags $466,183 $466,183

 Other Variable Costs $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969 $26,693,969

 Fixed Operating Costs $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826 $10,655,826

Total Product Cost $82,550,916 $82,550,916 $82,550,916 $82,550,916 $83,017,099 $82,550,916 $82,550,916 $82,550,916 $82,550,916 $83,017,099

Annual Depreciation

General Plant Writedown

 Depreciation Charge

 Remaining Value

Steam Plant Writedown 4.46% 4.46% 4.46% 4.46% 2.23%

 Depreciation Charge $2,943,116 $2,942,457 $2,943,116 $2,942,457 $1,471,558

 Remaining Value $10,299,587 $7,357,131 $4,414,015 $1,471,558 $0

Net Revenue $52,497,149 $52,497,808 $52,497,149 $52,497,808 $53,502,524 $55,440,265 $55,440,265 $55,440,265 $55,440,265 $54,974,082

Losses Forward $0 $0 $0 $0 $0 $0 $0 $0 $0 $0

Taxable Income $52,497,149 $52,497,808 $52,497,149 $52,497,808 $53,502,524 $55,440,265 $55,440,265 $55,440,265 $55,440,265 $54,974,082

Income Tax $18,374,002 $18,374,233 $18,374,002 $18,374,233 $18,725,883 $19,404,093 $19,404,093 $19,404,093 $19,404,093 $19,240,929

Annual Cash Income $37,066,263 $37,066,032 $37,066,263 $37,066,032 $36,248,199 $36,036,172 $36,036,172 $36,036,172 $36,036,172 $35,733,153

Discount Factor 0.1978 0.1799 0.1635 0.1486 0.1351 0.1228 0.1117 0.1015 0.0923 0.0839

Annual Present Value $7,333,362 $6,666,652 $6,060,630 $5,509,629 $4,898,240 $4,426,899 $4,024,453 $3,658,594 $3,325,994 $2,998,206

Total Capital Investment + Interest

Net Present Worth

Year 27 28 29 30

Fixed Capital Investment

Land ($1,848,000)

Working Capital ($20,030,673)

Loan Payment $0 $0 $0 $0

 Loan Interest Payment $0 $0 $0 $0

 Loan Principal $0 $0 $0 $0

 Ethanol Sales $131,425,440 $131,425,440 $131,425,440 $131,425,440

 By-Product Credit $6,565,742 $6,565,742 $6,565,742 $6,565,742

Total Annual Sales $137,991,181 $137,991,181 $137,991,181 $137,991,181

Annual Manufacturing Cost

 Feedstock $45,201,122 $45,201,122 $45,201,122 $45,201,122

 Baghouse Bags

 Other Variable Costs $26,693,969 $26,693,969 $26,693,969 $26,693,969

 Fixed Operating Costs $10,655,826 $10,655,826 $10,655,826 $10,655,826

Total Product Cost $82,550,916 $82,550,916 $82,550,916 $82,550,916

Annual Depreciation

General Plant Writedown

 Depreciation Charge

 Remaining Value

Steam Plant Writedown

 Depreciation Charge

 Remaining Value

Net Revenue $55,440,265 $55,440,265 $55,440,265 $55,440,265

Losses Forward $0 $0 $0 $0

Taxable Income $55,440,265 $55,440,265 $55,440,265 $55,440,265

Income Tax $19,404,093 $19,404,093 $19,404,093 $19,404,093

Annual Cash Income $36,036,172 $36,036,172 $36,036,172 $36,036,172

Discount Factor 0.0763 0.0693 0.0630 0.1486

Annual Present Value $2,748,756 $2,498,869 $2,271,699 $5,356,547

Total Capital Investment + Interest ($3,252,125)

Net Present Worth
106 

Appendix C. Process Parameters/Operating Summary

Feedstock Conditioning Boiler/Turbogenerator

Moisture Content 20% Ammonia Loading (% stoich. XS 0.0% WWT sludge moisture % 79.0%

Pretreatment Ammonia Loading (g/L hydrolyzate) 4.7 First effect lignin moisture % 35.0%

Acid Conc (wt%) 0.82 Sugar Losses: Combined boiler feed moisture 44.3%

Acid Loading (mg/g dry biomass) 22.1 Xylose 1.0% Boiler feed combustion energy MMkcal/hr -181

Total Solids (wt%) 30.0% Arabinose 1.0% % from lignin cake 54%

Insoluble Solids in (wt%) 22.6% Glucose 0.0% % from biogas 36%

Insoluble Solids out (wt%) 16.6% Galactose 1.0% Boiler feed heating value (kcal/kg) -497

Temperature (°C) 158 Mannose 1.0% Boiler feed heating value (Btu/lb) -895

Pressure (atm) 5.7 Cellobiose 1.0%

Conversions Overall Fermentation Boiler efficiency (feed/steam) 80%

 Cellulose to Glucolig 0.3% Total Solids (wt%) 19.8%

 Cellulose to Cellobiose 0.0% Insoluble Solids (wt%) 5.1% Steam loop flow (kg/hr) 234,784

 Cellulose to Glucose 9.9% Temperature (°C) 32 Total Generator Output (kW) -41,324

 Cellulose to HMF 0.3% Pressure (atm) 1.0 to A100 Feedstock Handling 859

 Xylan to Oligomer 2.4% Residence Time (days) 1.5 to A200 Pretreatment 5,680

 Xylan to Xylose 90.0% Conversions: to A300 Hydrolysis & Fermentation 2,636

 Xylan to Furfural 5.0% Glucose to Ethanol 95.0% to A400 Enzyme Production 5,340

 Xylan to Tar 0.0% Glucose to Zymo (cell mass) 2.0% to A500 Recovery 2,124

 Mannan to Oligomer 2.4% Glucose to Glycerol 0.4% to A600 Wastewater 7,368

 Mannan to Mannose 90.0% Glucose to Succinic Acid 0.6% to A700 Storage 11

 Mannan to HMF 5.0% Glucose to Acetic Acid 0.0% to A800 Boiler/Turbogen 1,368

 Galactan to Oligomer 2.4% Glucose to Lactic Acid 0.0% to A900 Utilities 3,141

 Galactan to Galactose 90.0% Xylose to Ethanol 85.0% Excess electricity to grid 12,797

 Galactan to HMF 5.0% Xylose to Zymo 1.9% Electricity Used (kW) 28,527

 Arabinan to Oligomer 2.4% Xylose to Glycerol 0.3% Electricity Used (%) 69%

 Arabinan to Arabinose 90.0% Xylose to Xylitol 4.6% Wastewater Treatment

 Arabinan to Furfural 5.0% Xylose to Succinic Acid 0.9% NH3-N to Ferm mg/L 1,608

 Arabinan to Tar 0.0% Xylose to Acetic Acid 0.0% NH3-N out of Ferm mg/L 1,565

 Acetate to Oligomer 0.0% Xylose to Lactic Acid 0.0% Protein-N 0

 Acetate to Acetic Acid 100.0% Arabinose to Ethanol 85.0% Cellulase-N 202

 Furfural to Tar 100.0% Arabinose to Zymo 1.9% NH3-N in Combined WW mg/L 667

 HMF to Tar 100.0% Arabinose to Glycerol 0.3% Protein-N 0

 Lignin to Soluble Lignin 5.0% Arabinose to Succinic Acid 1.5% Cellulase-N 43

 Xylose olig to xylose Arabinose to Acetic Acid 0.0% Hydraulic Load (L/hr) 416,673

 Mannose olig to mannose Arabinose to Lactic Acid 0.0% Hydraulic Load (gpm) 1,835

 Arabinose olig to arabinose Galactose to Ethanol 0.0% Hydraulic Load (MMgal/day) 2.6

 Galactose olig to galactose Galactose to Zymo 0.0% Total COD (kg/hr) 26,536

Enzymatic Hydrolysis Galactose to Glycerol 0.0% Total COD (g/L) 63.7

Enzyme Loading (mg/g cell) 19.9 Galactose to Succinic Acid 0.0%

Total Solids (wt%) 20.0% Galactose to Acetic Acid 0.0% Biogas composition (Dry molar%) CH4 51%

Insoluble Solids (wt%) 10.6% Galactose to Lactic Acid 0.0% CO2 49%

Temperature (°C) 32 Mannose to Ethanol 0.0% methane produced (kg/kg COD) 0.223

Pressure (atm) 1.0 Mannose to Zymo 0.0% cell mass 0.044

Residence Time (days) 3.5 Mannose to Glycerol 0.0%

Conversions: Overall Mannose to Succinic Acid 0.0%

 Cellulose to Glucolig 4.0% Mannose to Acetic Acid 0.0%

 Cellulose to Cellobiose 1.2% Mannose to Lactic Acid 0.0%

 Cellulose to Glucose 90.0%

 Glucolig to Cellobiose 0.0% Corn Steep Liquor loading wt% 0.00%

 Glucolig to Glucose 0.0% Contamination Loss 3.0%

 Cellobiose to Glucose 0.0% Ethanol Out of Fermenters (wt%) 5.4%

 Xylan to Oligomer 0.0% Enzyme Production

 Xylan to Xylose 0.0% Glucose feed (kg/hr) 2,418

 Xylose Oligomer to Xylose 0.0% Protein produced (kg/hr) 579

 Xylan to Tar 0.0% Protein yield (kg/kg substrate) 0.24

 Arabinan to Oligomer 0.0% Protein harvest titer (g/L) 40

 Arabinan to Arabinose 0.0%

 Galactan to Oligomer 0.0% Compressor electricity (kW) 1408

 Galactan to Galactose 0.0% Agitator electricity (kW) 2093

 Galactose Oligomer to Galactose 0.0% Chiller system electricity (kW) 1587

 Mannan to Oligomer 0.0% A400 total electricity (kW) 5340

 Mannan to Mannose 0.0%

 Mannose Oligomer to Mannose 0.0% Electricity demand (kWh/kg protn) 9

Production selectivity:

 Selectivity to protein 31.3%

 Selectivity to cell mass 4.1%

 Selectivity to CO2 67.6%
107 

Appendix D. Aspen Properties

The first table in this appendix is a list of the components used in the Aspen model. Previous 

versions of the model used custom property databanks created at NREL. Where possible in the 

new model, these components have been replaced with components from Aspen’s native 

databanks. Property definitions for the few remaining custom components were moved into the 

model itself (i.e., inside the simulation file) and are discussed here.

The second table is a list of combustion energies (LHV and HHV) computed using heats of 

formation from Aspen. 

The component CSL (corn steep liquor) used in previous models was removed. CSL streams are 

now modeled as 50% water, 25% protein, and 25% lactic acid [89].
108 

Component Property Quantity Units Reference

H2O - - - Native Aspen component

ETHANOL - - - Native Aspen component

GLUCOSE - - - Native Aspen component (dextrose)

GALACTOS - - - Duplicate of GLUCOSE

MANNOSE - - - Duplicate of GLUCOSE

XYLOSE DHFORM -216752.65 cal/mol Native Aspen component (d-xylose) with DHFORM specified (5/6 of GLUCOSE 

DHFORM)

ARABINOS - - - Duplicate of XYLOSE

CELLOB - - - Cellobiose. Used native Aspen component sucrose

SUCROSE - - - Native Aspen component

GLUCOLIG MW 162.1424 Glucose oligomers. Most properties from GLUCOSE; MW is GLUCOSE minus 

H2O

DHFORM -192875.34 cal/mol Back-calculated to match ΔHc of CELLULOS

GALAOLIG - - - Galactose oligomers. Duplicate of GLUCOLIG

MANOLIG - - - Mannose oligomers. Duplicate of GLUCOLIG

XYLOLIG MW 132.11612 Xylose oligomers. Most properties from XYLOSE; MW is XYLOSE minus H2O

DHFORM -149412.58 cal/mol Back-calculated to match ΔHc of XYLAN

ARABOLIG - - - Arabinose oligomers. Duplicate of XYLOLIG

EXTRACT - - - Organic extractives. Duplicate of GLUCOSE

LGNSOL - - - Solubilized lignin. Native Aspen component vanillin (see note at LIGNIN)
109 

Component Property Quantity Units Reference

HMF MW 126.11 (5-hydroxymethylfurfural) Properties for HMF were estimated within Aspen using 

NIST TDE routines. Specify molecular structure, MW, TB and DHFORM. From 

[90].

TB 532.7 K

DHFORM -79774.53 cal/mol

DHVLWT180550000 J/kmol

TC 731.012 K NIST TDE

PC 5235810 Pa

OMEGA 0.993646710

VC 0.3425 m3/kmol

RKTZRA 0.198177974

FURFURAL - - - Native Aspen component

AACID - - - Native Aspen component (acetic acid)

LACID - - - Native Aspen component (lactic acid)

XYLITOL - - - Native Aspen component

GLYCEROL - - - Native Aspen component

SUCCACID - - - Native Aspen component (Succinic acid)

NH3 - - - Native Aspen component

H2SO4 - - - Native Aspen component

NH4SO4 - - - Native Aspen component (ammonium sulfate)

NH4ACET PLXANT/1 -1E20 atm Native Aspen component (ammonium acetate) forced non-volatile

DAP - - - Native Aspen component (diammonium phosphate)

HNO3 - - - Native Aspen component

NANO3 - - - Native Aspen component

NAOH - - - Native Aspen component

CNUTR - - - Cellulase nutrient mix. Duplicate of glucose

WNUTR - - - WWT nutrient mix. Duplicate of glucose

DENAT - - - Denaturant. Native Aspen component (n-heptane)

OIL - - - Corn oil antifoam. Native Aspen component (oleic acid)

O2 - - - Native Aspen component
110 

Component Property Quantity Units Reference

N2 - - - Native Aspen component

NO - - - Native Aspen component

NO2 - - - Native Aspen component

CO - - - Native Aspen component

CO2 - - - Native Aspen component

CH4 - - - Native Aspen component

H2S - - - Native Aspen component

SO2 - - - Native Aspen component

CELLULOS DHSFRM -233200.06 cal/mol Native Aspen component with specified heat of formation; back-calculated 

GALACTAN - - - Duplicate of CELLULOS

MANNAN - - - Duplicate of CELLULOS

XYLAN Formula C5H8O4 (monomer)

MW 132.117

DHSFRM -182099.93 cal/mol From [22]. Assumes the ratio of ΔHc of glucose:xylose is the same for 

cellulose:xylan.

ARABINAN - - - Duplicate of xylan

LIGNIN - - - Used native Aspen component vanillin (C8H8O3). The HHV of this compound (-

23,906 BTU/kg) is very close to what we previously assumed for lignin as a 

custom component (-24,206)

ACETATE - - - Used native Aspen component acetic acid

PROTEIN Formula CH1.57O0.31N0.29S0.007 Wheat gliadin [91]

MW 22.8396

DHSFRM -17618 cal/mol From literature value of gliadin ΔHc [92] 

PLXANT/1 -1E20 atm Forces non-volatility

ASH - - - Native Aspen component CaO

ENZYME Formula CH1.59O0.42N0.24S0.01 Provided by Novozymes [55]

MW 24.0156

DHSFRM -17618 cal/mol copied from PROTEIN

DENZ - - - Denatured enzyme. Duplicate of enzyme
111 

Component Property Quantity Units Reference

ZYMO Formula CH1.8O0.5N0.2 Z. mobilis cell mass. Average composition of several microorganisms outlined

in [93]

MW 24.6264

DHSFRM -31169.39 cal/mol From [22] after [94]

TRICHO Formula CH1.645O0.445N0.205S0.005 T. reesei cell mass. Average of generic cell mass [56] and the ENZYME 

composition

MW 23.8204

DHSFRM -23200.01 cal/mol Copied BIOMASS

BIOMASS Formula CH1.64O0.39N0.23S0.0035 WWT sludge.

MW 23.238

DHSFRM -23200.01 cal/mol From [22] after [94] 

TAR - - - Modeled as solid xylose

LIME - - - Aspen native component (calcium hydroxide)

CASO4 - - - Aspen native component

C - - - Aspen native component (graphite carbon)
112 

MW ΔHf (IG) ΔHf (L) ΔHf (S)

cal/mol cal/mol cal/mol cal/mol Gcal/kmol BTU/lb cal/mol Gcal/kmol BTU/lb

GLUCOSE 180.16 -260103 -300428 GLUCOSE + 6 O2 → 6 H2O + 6 CO2 -610039 -0.610039 -6095 -672895 -0.672895 -6723

XYLOSE 150.13 -216753 -249440 XYLOSE + 5 O2 → 5 H2O + 5 CO2 -509283 -0.509283 -6106 -561662 -0.561662 -6734

SUCROSE 342.30 -459301 -480900 SUCROSE + 12 O2 → 11 H2O + 12 CO2 -1282278 -1.282278 -6743 -1397513 -1.397513 -7349

EXTRACT 180.16 -260103 -300428 EXTRACT + 6 O2 → 6 H2O + 6 CO2 -610039 -0.610039 -6095 -672895 -0.672895 -6723

ARABINOS 150.13 -216753 -249440 ARABINOS + 5 O2 → 5 H2O + 5 CO2 -509283 -0.509283 -6106 -561662 -0.561662 -6734

GALACTOS 180.16 -260103 -300428 GALACTOS + 6 O2 → 6 H2O + 6 CO2 -610039 -0.610039 -6095 -672895 -0.672895 -6723

MANNOSE 180.16 -260103 -300428 MANNOSE + 6 O2 → 6 H2O + 6 CO2 -610039 -0.610039 -6095 -672895 -0.672895 -6723

GLUCOLIG 162.14 -192875 -233200 GLUCOLIG + 6 O2 → 5 H2O + 6 CO2 -619511 -0.619511 -6877 -671890 -0.671890 -7459

CELLOB 342.30 -459301 -480900 CELLOB + 12 O2 → 11 H2O + 12 CO2 -1282278 -1.282278 -6743 -1397513 -1.397513 -7349

XYLOLIG 132.12 -149413 -182100 XYLOLIG + 5 O2 → 4 H2O + 5 CO2 -518866 -0.518866 -7069 -560770 -0.560770 -7640

MANOLIG 162.14 -192875 -233200 MANOLIG + 6 O2 → 5 H2O + 6 CO2 -619511 -0.619511 -6877 -671890 -0.671890 -7459

GALAOLIG 162.14 -192875 -233200 GALAOLIG + 6 O2 → 5 H2O + 6 CO2 -619511 -0.619511 -6877 -671890 -0.671890 -7459

ARABOLIG 132.12 -149413 -182100 ARABOLIG + 5 O2 → 4 H2O + 5 CO2 -518866 -0.518866 -7069 -560770 -0.560770 -7640

XYLITOL 152.15 -222246 -243145 -267149 XYLITOL + 5.5 O2 → 6 H2O + 5 CO2 -573334 -0.573334 -6783 -636190 -0.636190 -7527

LGNSOL 152.15 -88134 -108248 LGNSOL + 8.5 O2 → 4 H2O + 8 CO2 -874683 -0.874683 -10348 -916587 -0.916587 -10844

ETHANOL 46.07 -56117 -66278 ETHANOL + 3 O2 → 3 H2O + 2 CO2 -294968 -0.294968 -11525 -326395 -0.326395 -12753

H2O 18.02 -57756 -68232 -69963 NA

FURFURAL 96.09 -36066 -47738 FURFURAL + 5 O2 → 2 H2O + 5 CO2 -537716 -0.537716 -10073 -558667 -0.558667 -10466

HMF 126.11 -79775 -99677 HMF + 6 O2 → 3 H2O + 6 CO2 -637521 -0.637521 -9099 -668949 -0.668949 -9548

H2SO4 98.08 -175600 -189490 H2SO4 + -0.5 O2 → 1 H2O + 1 SO2 60835 0.060835 1116 50359 0.050359 924

N2 28.01 0 -- -- NA

CO2 44.01 -93988 -95245 NA

O2 32.00 0 -- -- NA

CH4 16.04 -17799 -- -- CH4 + 2 O2 → 2 H2O + 1 CO2 -191702 -0.191702 -21509 -212654 -0.212654 -23860

NO 30.01 21556 -- NO + 0.5 O2 → 1 NO2 -13631 -0.013631 -818 -13631 -0.013631 -818

NO2 46.01 7925 -1179 NA

NH3 17.03 -10963 -15706 NH3 + 0.75 O2 → 1.5 H2O + 0.5 N2 -75672 -0.075672 -7998 -91386 -0.091386 -9659

LACID 90.08 -143212 -163122 LACID + 3 O2 → 3 H2O + 3 CO2 -292112 -0.292112 -5837 -323539 -0.323539 -6465

AACID 60.05 -103373 -108966 AACID + 2 O2 → 2 H2O + 2 CO2 -194523 -0.194523 -5831 -215474 -0.215474 -6459

NH4SO4 132.14 -282029 -288994 -282041 NH4SO4 + 1 O2 → 4 H2O + 1 N2 + 1 SO2 -12930 -0.012930 -176 -54834 -0.054834 -747

NH4ACET 77.08 -147163 -154701 -147163 NH4ACET + 2.75 O2 → 3.5 H2O + 2 CO2 + 0.5 N2 -235422 -0.235422 -5498 -272088 -0.272088 -6354

GLYCEROL 92.09 -138029 -159483 GLYCEROL + 3.5 O2 → 4 H2O + 3 CO2 -353507 -0.353507 -6909 -395411 -0.395411 -7728

SUCCACID 118.09 -196546 -221160 -231172 SUCCACID + 3.5 O2 → 3 H2O + 4 CO2 -328062 -0.328062 -5001 -359490 -0.359490 -5480

DAP 132.06 -276457 -283996 -374250 NA

DENAT 100.20 -44819 -53568 DENAT + 11 O2 → 8 H2O + 7 CO2 -1066400 -1.066400 -19156 -1150207 -1.150207 -20662

OIL 282.47 -169628 -196067 OIL + 25.5 O2 → 17 H2O + 18 CO2 -2477579 -2.477579 -15788 -2655669 -2.655669 -16923

CNUTR 180.16 -260103 -300428 CNUTR + 6 O2 → 6 H2O + 6 CO2 -610039 -0.610039 -6095 -672895 -0.672895 -6723

WNUTR 180.16 -260103 -300428 WNUTR + 6 O2 → 6 H2O + 6 CO2 -610039 -0.610039 -6095 -672895 -0.672895 -6723

LHV HHV
113 

MW ΔHf (IG) ΔHf (L) ΔHf (S)

cal/mol cal/mol cal/mol cal/mol Gcal/kmol BTU/lb cal/mol Gcal/kmol BTU/lb

SO2 64.06 -70899 -76380 NA

H2S 34.08 -4927 -8300 H2S + 1.5 O2 → 1 H2O + 1 SO2 -123728 -0.123728 -6535 -134204 -0.134204 -7088

CO 28.01 -26400 -- CO + 0.5 O2 → 1 CO2 -67589 -0.067589 -4343 -67589 -0.067589 -4343

HNO3 63.01 -32077 -41406 HNO3 + -1.75 O2 → 0.5 H2O + 0.5 N2 12528 0.012528 358 -2039 -0.002039 -58

NA2O -6965 NA

NAOH 40.00 -47234 -67046 NAOH + 0 O2 → 0.5 H2O + 0.5 NA2O 34685 0.034685 1561 9636 0.009636 434

NANO3 84.99 -118756 NANO3 + -1.25 O2 → 0.5 N2 + 0.5 NA2O 115273 0.115273 2441 115273 0.115273 2441

CELLULOS 162.14 -233200 CELLULOS + 6 O2 → 5 H2O + 6 CO2 -619511 -0.619511 -6877 -671890 -0.671890 -7459

XYLAN 132.12 -182100 XYLAN + 5 O2 → 4 H2O + 5 CO2 -518866 -0.518866 -7069 -560770 -0.560770 -7640

LIGNIN 152.15 -108248 LIGNIN + 8.5 O2 → 4 H2O + 8 CO2 -874683 -0.874683 -10348 -916587 -0.916587 -10844

ENZYME 24.02 -17900 ENZYME + 1.1975 O2 → 0.795 H2O + 1 CO2 + 0.12 N2 + 0.01 SO2 -122713 -0.122713 -9198 -130333 -0.130333 -9769

DENZ 24.02 -17900 DENZ + 1.1975 O2 → 0.795 H2O + 1 CO2 + 0.12 N2 + 0.01 SO2 -122713 -0.122713 -9198 -130333 -0.130333 -9769

BIOMASS 23.24 -23200 BIOMASS + 1.2185 O2 → 0.82 H2O + 1 CO2 + 0.115 N2 + 0.0035 SO2 -118397 -0.118397 -9171 -126739 -0.126739 -9817

ZYMO 24.63 -31169 ZYMO + 1.2 O2 → 0.9 H2O + 1 CO2 + 0.1 N2 -114800 -0.114800 -8391 -124228 -0.124228 -9080

ACETATE 60.05 -103373 -108992 ACETATE + 2 O2 → 2 H2O + 2 CO2 -194497 -0.194497 -5830 -215449 -0.215449 -6458

ARABINAN 132.12 -182100 ARABINAN + 5 O2 → 4 H2O + 5 CO2 -518866 -0.518866 -7069 -560770 -0.560770 -7640

MANNAN 162.14 -233200 MANNAN + 6 O2 → 5 H2O + 6 CO2 -619511 -0.619511 -6877 -671890 -0.671890 -7459

GALACTAN 162.14 -233200 GALACTAN + 6 O2 → 5 H2O + 6 CO2 -619511 -0.619511 -6877 -671890 -0.671890 -7459

TAR 150.13 -182100 TAR + 5 O2 → 5 H2O + 5 CO2 -576623 -0.576623 -6913 -629002 -0.629002 -7541

ASH 56.08 10485 -151688 NA

TRICHO 23.82 -23200 TRICHO + 1.19375 O2 → 0.8225 H2O + 1 CO2 + 0.1025 N2 + 0.005 SO2 -118647 -0.118647 -8966 -126909 -0.126909 -9590

PROTEIN 22.84 -17618 PROTEIN + 1.2445 O2 → 0.785 H2O + 1 CO2 + 0.145 N2 + 0.007 SO2 -122205 -0.122205 -9631 -129932 -0.129932 -10240

CARBON 12.01 0 C + 1 O2 → 1 CO2 -93988 -0.093988 -14085 -93988 -0.093988 -14085

LIME 74.09 -145878 -235522 Ca(OH)2 → 1 H2O + 1 Ash 26078 0.026078 634 15602 0.015602 379

CASO4 136.14 -315682 -342531 CaSO4 + -0.5 O2 → 1 SO2 + 1 Ash 119944 0.119944 1586 119944 0.119944 1586

LHV HHV
114 

Appendix E. Process Flow Diagrams

The full set of PFDs for this process follow. Space for stream tables was limited; below is a key 

to lumped components. See Section 5.1 for a discussion on how LHV is computed. 

Other sugars (SS) arabinose, mannose, galactose, cellobiose, sucrose

Sugar oligomers (SS) oligomers of glucose, xylose, and all minor sugars

Organic soluble solids (SS) ammonium acetate, solubilized lignin, organic extractives, lactic acid, 

cellulase and WWT nutrients

Inorganic soluble solids (SS) ammonium sulfate, diammonium phosphate, sodium hydroxide, nitric 

acid, sodium nitrate

Furfurals furfural, HMF

Other organics glycerol, gasoline denaturant, corn oil, succinic acid, xylitol

CO/SOX/NOX/H2S NO, NO2, SO2, CO, H2S

Other struct. carbohydr. (IS) arabinan, mannan, galactan

Protein (IS) corn protein, cellulase, denatured cellulase

Cell mass (IS) WWT sludge, Z. mobilis from fermentation, T. reesei from cellulase 

production

Other insoluble solids (IS) tar, ash, carbon, lime, gypsum
Work Stream kW Gcal/hr

WC101 32 0.027

WC102 79 0.068

WC103 16 0.014

WC104 96 0.082

WC105 16 0.014

WC106 16 0.014

WC1078 0.007

WC1084 0.003

WM102 317 0.272

WM103 158 0.136

WM106 119 0.102

COMPONENT UNITS 101 105 840

Total Flow kg/hr 104,167 104,167 0

Insoluble Solids % 67.7% 67.7% 0.0%

Soluble Solids % 12.3% 12.3% 0.0%

Temperature °C 25 25 25

Pressure atm 1.0 1.0 1.0

Vapor Fraction 0.0 0.0 1.0

Ethanol kg/hr

Water kg/hr 20,833 20,833

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr 642 642

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr 12,208 12,208

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr

Carbon Dioxide kg/hr

Methane kg/hr0

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr 29,205 29,205

Xylan (IS) kg/hr 16,273 16,273

Other Struct. Carbohydr. (IS) kg/hr 3,675 3,675

Acetate (IS) kg/hr 1,508 1,508

Lignin (IS) kg/hr 13,132 13,132

Protein (IS) kg/hr 2,583 2,583

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr 4,108 4,108

Carbon Mole Flow kmol/hr 3,117 3,117 0

COD kg/hr 101,238 101,238 0

LHV | 25°C Gcal/hr -315.6 -315.6 0.0

115
COMPONENT UNITS 105 212 214 215 216 217 218 222 232 516 710

Total Flow kg/hr 104,167 140,850 38,801 3,490 24,534 311,842 19,782 292,060 1,980 36,821 1,981

Insoluble Solids % 67.7% 0.0% 0.0% 0.0% 0.0% 14.8% 0.0% 15.8% 0.0% 0.0% 0.0%

Soluble Solids % 12.3% 0.0% 0.0% 0.0% 0.0% 12.0% 0.0% 12.8% 0.0% 0.0% 0.0%

Temperature °C 25 95 113 268 268 158 130 130 22 114 21

Pressure atm 1.0 4.7 6.1 13.0 13.0 5.7 2.6 9.6 11.8 6.1 5.4

Vapor Fraction 0.0 0.0 0.0 1.0 1.0 0.0 1.0 0.0 0.0 0.0 0.0

Ethanol kg/hr 18 18 8 10 18

Water kg/hr 20,833 140,850 36,767 3,490 24,534 224,002 19,533 204,469 139 36,629 139

Glucose (SS) kg/hr0 3,550 0 3,5500

Xylose (SS) kg/hr0 16,643 0 16,6430

Other sugars (SS) kg/hr 6420 3,720 0 3,7200

Sugar Oligomers (SS) kg/hr0 566 0 5660

Organic Soluble Solids (SS) kg/hr 12,2081 12,866 0 12,8661

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr00000

Acetic Acid kg/hr 1,508 36 1,473

Sulfuric Acid kg/hr 1,842 1,842 0 1,842 1,842 1,842

Furfurals kg/hr 172 1,034 205 829 172

Other Organics kg/hr00000

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr00

Cellulose (IS) kg/hr 29,205 26,138 26,138

Xylan (IS) kg/hr 16,273 423 423

Other Struct. Carbohydr. (IS) kg/hr 3,675 96 96

Acetate (IS) kg/hr 1,508

Lignin (IS) kg/hr 13,132 12,475 12,475

Protein (IS) kg/hr 2,583 2,583 2,583

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr 4,108 4,377 4,377

Carbon Mole Flow kmol/hr 3,117 10 3,127 12 3,115 10

COD kg/hr 101,238 327 101,565 396 101,168 327

LHV | 25°C Gcal/hr -315.6 72.2 18.2 -0.4 -2.7 -231.7 -2.9 -230.0 1.2 16.9 1.2

Work Stream kW Gcal/hr

WC201 80 0.069

WC202 80 0.069

WC203 160 0.138

WC204 75 0.064

WC205 80 0.069

WC206 20 0.017

WC207 40 0.034

WM201 80 0.069

WM202 30 0.026

WM203 2000 1.720

WM204 45 0.039

WM205 30 0.026

WM206 2000 1.720

WM207 82 0.071

WP2011 0.001

WP203 64 0.055

WT203 170 0.146

116
COMPONENT UNITS 211 212 222 223 224 250 251 252 253 254 255 260 273 274 275 301

Total Flow kg/hr 140,850 140,850 292,060 292,060 13,866 15,932 3,850 13,682 16,117 278,194 13,682 33,648 1,051 150,310 151,360 429,554

Insoluble Solids % 0.0% 0.0% 15.8% 15.8% 0.0% 0.0% 0.0% 0.0% 0.0% 16.6% 0.0% 0.0% 0.0% 0.0% 0.0% 10.8%

Soluble Solids % 0.0% 0.0% 12.8% 12.8% 0.0% 0.0% 0.0% 0.0% 0.0% 13.4% 0.0% 0.0% 0.0% 0.0% 0.0% 9.7%

Temperature °C 33 95 130 131 101 100 100 100 100 97 99 100 20 33 36 78

Pressure atm 5.0 4.7 9.6 5.7 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 9.0 5.0 5.0 4.3

Vapor Fraction 0.0 0.0 0.0 0.0 1.0 1.0 0.0 1.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

Ethanol kg/hr 10 10 4 8 0 11 1 6 11 136

Water kg/hr 140,850 140,850 204,469 204,469 13,715 15,702 3,831 13,382 16,035 190,754 13,382 33,248 150,310 150,310 341,063

Glucose (SS) kg/hr 3,550 3,550 0 0 0 0 0 3,550 00 3,550

Xylose (SS) kg/hr 16,643 16,643 0 0 0 0 0 16,643 00 16,476

Other sugars (SS) kg/hr 3,720 3,720 0 0 0 0 0 3,720 00 3,682

Sugar Oligomers (SS) kg/hr 566 566 0 0 0 0 0 566 00 566

Organic Soluble Solids (SS) kg/hr 12,866 12,866 0 0 0 0 0 12,866 00 14,727

Inorganic Soluble Solids (SS) kg/hr 2,482

Ammonia kg/hr0000000000 1,051 1,051

Acetic Acid kg/hr 1,473 1,473 23 25 11 15 32 1,450 15 58

Sulfuric Acid kg/hr 1,842 1,842 0 0 0 0 0 1,842 00

Furfurals kg/hr 829 829 124 197 8 274 48 705 274 329 705

Other Organics kg/hr00000000000

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr 26,138 26,138 26,138 26,138

Xylan (IS) kg/hr 423 423 423 423

Other Struct. Carbohydr. (IS) kg/hr 96 96 96 96

Acetate (IS) kg/hr

Lignin (IS) kg/hr 12,475 12,475 12,475 12,475

Protein (IS) kg/hr 2,583 2,583 2,583 2,583

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr 4,377 4,377 4,377 4,581

Carbon Mole Flow kmol/hr 3,115 3,115 7 11 1 15 4 3,107 15 20 3,107

COD kg/hr 101,168 101,168 239 372 25 495 116 100,929 495 636 100,929

LHV | 25°C Gcal/hr 80.9 72.2 -230.0 -230.2 -1.5 -1.7 1.9 -2.1 7.8 -228.9 5.1 14.8 -4.4 86.3 81.3 -147.2

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QH201 -8.636 WP204 56 0.048

QH244 7.276 WP208 100 0.086

WP209 90 0.077

WT204 127 0.109

WT208 127 0.109

WT209 90 0.077

117
COMPONENT UNITS 303 304 305 309 310

Total Flow kg/hr 44,339 42,607 1,969 211 26

Insoluble Solids % 5.0% 5.7% 0.0% 25.0% 0.0%

Soluble Solids % 15.6% 6.3% 0.0% 24.8% 100.0%

Temperature °C 32 33 32 20 20

Pressure atm 1.0 5.0 1.0 1.0 1.0

Vapor Fraction 0.0 0.0 1.0 0.0 0.0

Ethanol kg/hr 1 2,009 32

Water kg/hr 35,149 35,248 38 106

Glucose (SS) kg/hr 2,969 148 0

Xylose (SS) kg/hr 1,648 168 0

Other sugars (SS) kg/hr 401 401 0

Sugar Oligomers (SS) kg/hr 161 161 0

Organic Soluble Solids (SS) kg/hr 1,477 1,529 0 52

Inorganic Soluble Solids (SS) kg/hr 248 272 26

Ammonia kg/hr 000

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr 71 70 0

Other Organics kg/hr 1 138 0

Carbon Dioxide kg/hr 0 51 1,882

Methane kg/hr

O2 kg/hr 00 17

N2 kg/hr 000

CO/SOX/NOX/H2S kg/hr 000

Cellulose (IS) kg/hr 125 125

Xylan (IS) kg/hr 42 42

Other Struct. Carbohydr. (IS) kg/hr 10 10

Acetate (IS) kg/hr

Lignin (IS) kg/hr 1,248 1,248

Protein (IS) kg/hr 320 364 53

Cell Mass (IS) kg/hr 11 162

Other Insoluble Solids (IS) kg/hr 458 458

Carbon Mole Flow kmol/hr 314 273 44 4

COD kg/hr 10,201 10,280 67 132

LHV | 25°C Gcal/hr -12.7 -12.6 -0.2 -0.4 0.0

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QF301-5 0.451 WF301-5 99 0.085

WP301 7 0.006

WP302 7 0.006

WT301 15 0.013

118
COMPONENT UNITS 301 302 303 304 306 311 312 315 317 422 501 551

Total Flow kg/hr 429,554 429,554 44,339 42,607 443,391 948 116 423,542 19,180 13,836 450,740 27,197

Insoluble Solids % 10.8% 10.8% 5.0% 5.7% 10.6% 25.0% 0.0% 5.5% 0.0% 5.3% 5.2% 0.0%

Soluble Solids% 9.7% 9.7% 15.6% 6.3% 9.4% 24.7% 100.0% 5.8% 0.0% 0.3% 5.4% 0.0%

Temperature °C 78 49 32 33 48 20 20 32 32 29 32 34

Pressure atm 4.3 2.5 1.0 5.0 5.1 1.0 1.0 1.0 1.0 1.9 6.0 4.3

Vapor Fraction 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.0 0.0 0.0 0.0

Ethanol kg/hr 661 2,009 6 21,517 331 21,941 424

Water kg/hr 341,063 341,063 35,149 35,248 354,117 476 351,782 369 13,054 378,525 26,743

Glucose (SS) kg/hr 3,550 3,550 2,969 148 3,550 521 0 521

Xylose (SS) kg/hr 16,476 16,476 1,648 168 16,476 1,062 0 1,062 0

Other sugars (SS) kg/hr 3,682 3,682 401 401 3,682 2,175 0 2,175 0

Sugar Oligomers (SS) kg/hr 566 566 161 161 566 1,612 0 1,612 0

Organic Soluble Solids (SS) kg/hr 14,727 14,727 1,477 1,529 14,768 235 16,421 0 41 16,421 0

Inorganic Soluble Solids (SS) kg/hr 2,482 2,482 248 272 2,482 116 2,610 2,610

Ammonia kg/hr00000000

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr 705 705 71 70 705 703 2 706 3

Other Organics kg/hr 001 138 13 1,400 0 13 1,400 0

Carbon Dioxide kg/hr0 51 1 520 18,323 1 549 28

Methane kg/hr

O2 kg/hr0000 155 000

N2 kg/hr00000000

CO/SOX/NOX/H2S kg/hr00000000

Cellulose (IS) kg/hr 26,138 26,138 125 125 26,138 1,255 1,255

Xylan (IS) kg/hr 423 423 42 42 423 423 423

Other Struct. Carbohydr. (IS) kg/hr 96 96 10 10 96 96 96

Acetate (IS) kg/hr

Lignin (IS) kg/hr 12,475 12,475 1,248 1,248 12,475 12,475 12,475

Protein (IS) kg/hr 2,583 2,583 320 364 3,203 237 3,445 620 3,445

Cell Mass (IS) kg/hr 11 162 108 944 108 944

Other Insoluble Solids (IS) kg/hr 4,581 4,581 458 458 4,581 4,581 4,581

Carbon Mole Flow kmol/hr 3,107 3,107 314 273 3,140 18 2,685 431 33 2,704 19

COD kg/hr 100,929 100,929 10,201 10,280 102,010 592 102,130 693 1,081 103,018 888

LHV | 25°C Gcal/hr -147.2 -136.5 -12.7 -12.6 -133.1 -1.7 0.0 -123.0 -2.1 3.6 -110.5 12.6

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QC301 10.636 WA308 75 0.064

QT300 4.417 WP306 106 0.091

QC310 5.946 WP310 55 0.047

WT300 1354 1.164

WT306 15 0.013

119
COMPONENT UNITS 403 404 405 406 409 435 450 452 453 454

Total Flow kg/hr 790 806 770 1,586 32,583 32,583 31,031 1,552

Insoluble Solids % 0.5% 25.0% 0.0% 0.0% 5.3% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 15.3% 24.7% 0.0% 0.0% 1.9% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 62 20 28 28 28 28 25 40 40 40

Pressure atm 3.0 1.0 3.0 9.0 2.0 1.0 1.0 3.0 3.0 3.0

Vapor Fraction 0.0 0.0 1.0 1.0 0.0 1.0 1.0 1.0 1.0 1.0

Ethanol kg/hr

Water kg/hr 665 4 714 37 638 638 608 30

Glucose (SS) kg/hr 121 12 0

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr 0220

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr 0600

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr 000

Carbon Dioxide kg/hr 00 95

Methane kg/hr

O2 kg/hr 00 287 7,440 7,440 7,086 354

N2 kg/hr 00 1,167 24,504 24,504 23,337 1,167

CO/SOX/NOX/H2S kg/hr 0000

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr 327

Cell Mass (IS) kg/hr 1 34

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr 4022

COD kg/hr 135 50 73 0

LHV | 25°C Gcal/hr -0.1 0.0 0.0 0.0 0.2 0.0 0.0 0.0 0.0 0.0

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QM401 1.172 WM401 1408 1.211

QF401-3 0.468 WP401 0 0.000

WT401-3 232 0.200

120
COMPONENT UNITS 400 401 402 403 405 409 410 416 422 423 440 441 442 443 446 453

Total Flow kg/hr 11,419 2,845 15,801 790 0 770 15,011 224 13,836 31,801 67 109 16 157 13 31,031

Insoluble Solids % 0.0% 0.0% 0.5% 0.5% 0.0% 5.3% 0.5% 17.5% 5.3% 0.0% 0.0% 0.0% 0.0% 25.0% 0.0% 0.0%

Soluble Solids % 0.0% 85.0% 15.3% 15.3% 0.0% 1.9% 15.3% 47.4% 0.3% 0.0% 100.0% 0.0% 0.0% 24.7% 0.0% 0.0%

Temperature °C 33 28 62 62 28 28 62 34 29 28 20 28 28 20 20 40

Pressure atm 5.0 1.0 3.0 3.0 3.0 2.0 3.0 2.0 1.9 1.0 1.0 9.0 3.0 1.0 3.2 3.0

Vapor Fraction 0.0 0.0 0.0 0.0 1.0 0.0 0.0 0.0 0.0 1.0 0.0 1.0 1.0 0.0 0.0 1.0

Ethanol kg/hr

Water kg/hr 11,419 427 13,296 665 714 12,631 79 13,054 737 79 608

Glucose (SS) kg/hr 2,418 2,418 121 12 2,297

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr5024 106 41 0 67 39

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr000000 109

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr1001 13 0 13

Carbon Dioxide kg/hr00001 2,287

Methane kg/hr

O2 kg/hr00000 5,440 7,086

N2 kg/hr00000 23,337 23,337

CO/SOX/NOX/H2S kg/hr0000000 16

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr 69 3 7 65 39 620 39

Cell Mass (IS) kg/hr 12 1 34 11 108

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr 81 84 4 2 80 5 33 52 2 3 1

COD kg/hr 2,577 2,697 135 0 73 2,562 170 1,081 0 72 4 98 39

LHV | 25°C Gcal/hr 6.6 -7.9 -1.4 -0.1 0.0 0.2 -1.3 -0.5 3.6 0.0 -0.2 -0.5 0.0 -0.3 -0.1 0.0

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QF400 8.104 WP400 1 0.001

QT405 -0.414 WP405 2 0.002

WP406 0 0.000

WP410 1 0.001

WP420 0 0.000

WT400 2093 1.800

WT405 6 0.005

WT406 2 0.002

WT410 7 0.006

121
COMPONENT UNITS 501 506 507 508 509 510

Total Flow kg/hr 450,740 450,740 391,501 391,501 610 58,629

Insoluble Solids % 5.2% 5.2% 5.9% 5.9% 0.0% 0.0%

Soluble Solids % 5.4% 5.4% 6.2% 6.2% 0.0% 0.0%

Temperature °C 32 103 125 47 60 117

Pressure atm 6.0 6.0 6.3 6.3 2.0 2.1

Vapor Fraction 0.0 0.0 0.0 0.0 1.0 1.0

Ethanol kg/hr 21,941 21,941 184 184 66 21,691

Water kg/hr 378,525 378,525 341,765 341,765 23 36,737

Glucose (SS) kg/hr 521 521 521 521 00

Xylose (SS) kg/hr 1,062 1,062 1,062 1,062 00

Other sugars (SS) kg/hr 2,175 2,175 2,175 2,175 00

Sugar Oligomers (SS) kg/hr 1,612 1,612 1,612 1,612 00

Organic Soluble Solids (SS) kg/hr 16,421 16,421 16,420 16,420 01

Inorganic Soluble Solids (SS) kg/hr 2,610 2,610 2,610 2,610

Ammonia kg/hr 000000

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr 706 706 533 533 0 172

Other Organics kg/hr 1,400 1,400 1,400 1,400 00

Carbon Dioxide kg/hr 549 549 522 27

Methane kg/hr

O2 kg/hr 000000

N2 kg/hr 000000

CO/SOX/NOX/H2S kg/hr 000000

Cellulose (IS) kg/hr 1,255 1,255 1,255 1,255

Xylan (IS) kg/hr 423 423 423 423

Other Struct. Carbohydr. (IS) kg/hr 96 96 96 96

Acetate (IS) kg/hr

Lignin (IS) kg/hr 12,475 12,475 12,475 12,475

Protein (IS) kg/hr 3,445 3,445 3,445 3,445

Cell Mass (IS) kg/hr 944 944 944 944

Other Insoluble Solids (IS) kg/hr 4,581 4,581 4,581 4,581

Carbon Mole Flow kmol/hr 2,704 2,704 1,738 1,738 15 951

COD kg/hr 103,018 103,018 57,393 57,393 137 45,489

LHV | 25°C Gcal/hr -110.5 -139.9 -25.6 3.8 -0.4 -137.3

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QCD501 0.462 WP501 309 0.266

QH512 29.396 WP503 1 0.000

QRD501 -33.022 WP506 56 0.048

122
COMPONENT UNITS 305 317 509 510 511 516 521 523 524 550 551

Total Flow kg/hr 1,969 19,180 610 58,629 29,213 36,821 7,405 21,759 26,836 21,398 27,197

Insoluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 32 32 60 117 91 114 71 33 33 34 34

Pressure atm 1.0 1.0 2.0 2.1 1.6 6.1 1.7 1.0 5.0 0.9 4.3

Vapor Fraction 1.0 1.0 1.0 1.0 1.0 0.0 0.0 1.0 0.0 1.0 0.0

Ethanol kg/hr 32 331 66 21,691 27,022 18 5,349 428 4 424

Water kg/hr 38 369 23 36,737 2,164 36,629 2,056 429 26,836 522 26,743

Glucose (SS) kg/hr 000000

Xylose (SS) kg/hr 0000000

Other sugars (SS) kg/hr 0000000

Sugar Oligomers (SS) kg/hr 0000000

Organic Soluble Solids (SS) kg/hr 00011000

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr 000000000

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr 020 172 0 172303

Other Organics kg/hr 00000000

Carbon Dioxide kg/hr 1,882 18,323 522 27 27 20,727 20,699 28

Methane kg/hr

O2 kg/hr 17 155 0 00 172 172 0

N2 kg/hr 00000000

CO/SOX/NOX/H2S kg/hr 000000000

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr 44 431 15 951 1,174 10 232 490 471 19

COD kg/hr 67 693 137 45,489 56,309 327 11,147 897 9 888

LHV | 25°C Gcal/hr -0.2 -2.1 -0.4 -137.3 -167.8 16.9 -33.3 -2.7 15.4 -0.1 12.6

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QCD502 23.180 WH505 223 0.192

QRD502 -5.654 WP504 8 0.007

WP505 13 0.011

WP515 5 0.004

123
COMPONENT UNITS 511 511A 512 514 515 517 519 521

Total Flow kg/hr 29,213 29,213 7,405 21,808 21,808 7,405 21,808 7,405

Insoluble Solids% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 91 116 116 116 38 35 92 71

Pressure atm 1.6 1.6 1.6 1.6 1.0 2.5 1.4 1.7

Vapor Fraction 1.0 1.0 1.0 1.0 0.0 0.0 1.0 0.0

Ethanol kg/hr 27,022 27,022 5,349 21,673 21,673 5,349 21,673 5,349

Water kg/hr 2,164 2,164 2,056 108 108 2,056 108 2,056

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr 00000

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr 00000

Other Organics kg/hr

Carbon Dioxide kg/hr 27 27 27 27 27

Methane kg/hr

O2 kg/hr 00000

N2 kg/hr 00000

CO/SOX/NOX/H2S kg/hr 00000

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr 1,174 1,174 232 941 941 232 941 232

COD kg/hr 56,309 56,309 11,147 45,162 45,162 11,147 45,162 11,147

LHV | 25°C Gcal/hr -167.8 -168.1 -33.3 -134.7 -138.9 -33.1 -134.5 -33.3

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QH503 2.565 WM503 151 0.130

QH506 -0.301

QH507 0.214

QH509 5.187

124
COMPONENT UNITS 508 535 559 560 571 572

Total Flow kg/hr 391,501 355,024 12,105 12,044 36,538 355,024

Insoluble Solids% 5.9% 0.1% 0.0% 0.0% 62.3% 0.1%

Soluble Solids% 6.2% 6.6% 0.0% 0.0% 2.5% 6.6%

Temperature °C 47 47 80 55 55 47

Pressure atm 6.3 3.9 6.9 6.9 6.9 1.0

Vapor Fraction 0.0 0.0 1.0 1.0 0.0 0.0

Ethanol kg/hr 184 177 1 6 177

Water kg/hr 341,765 329,030 237 175 12,797 329,030

Glucose (SS) kg/hr 521 502 0 19 502

Xylose (SS) kg/hr 1,062 1,022 0 40 1,022

Other sugars (SS) kg/hr 2,175 2,094 0 81 2,094

Sugar Oligomers (SS) kg/hr 1,612 1,552 0 60 1,552

Organic Soluble Solids (SS) kg/hr 16,420 15,808 0 612 15,808

Inorganic Soluble Solids (SS) kg/hr 2,610 2,513 97 2,513

Ammonia kg/hr 00000

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr 533 513 1 19 513

Other Organics kg/hr 1,400 1,348 0 52 1,348

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr 00 2,764 2,764 1 0

N2 kg/hr 00 9,104 9,103 1 0

CO/SOX/NOX/H2S kg/hr 000

Cellulose (IS) kg/hr 1,255 25 1,230 25

Xylan (IS) kg/hr 423 8 415 8

Other Struct. Carbohydr. (IS) kg/hr 96 2 94 2

Acetate (IS) kg/hr

Lignin (IS) kg/hr 12,475 250 12,226 250

Protein (IS) kg/hr 3,445 69 3,376 69

Cell Mass (IS) kg/hr 944 19 925 19

Other Insoluble Solids (IS) kg/hr 4,581 92 4,489 92

Carbon Mole Flow kmol/hr 1,738 801 0 937 801

COD kg/hr 57,393 25,900 4 31,489 25,900

LHV | 25°C Gcal/hr 3.8 100.8 -0.2 -0.1 -97.3 100.9

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QM505 0.093 WC501 7 0.006

QM507 0.733 WC502 11 0.009

WM505 112 0.096

WM507 1043 0.897

WP530 37 0.032

WP531 76 0.066

WP532 7 0.006

WP533 11 0.010

WP581 37 0.032

WT530 6 0.005

125
COMPONENT UNITS 260 535 601 602 611 612 615 640 821 903 904 944

Total Flow kg/hr 33,648 355,024 411,178 411,178 360,712 28,626 21,860 21 4,681 147,140 147,140 17,680

Insoluble Solids % 0.0% 0.1% 0.1% 0.1% 0.1% 4.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 6.6% 5.7% 5.7% 0.8% 0.8% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 100 47 52 35 35 35 38 86 100 13 59 28

Pressure atm 1.0 3.9 1.3 1.3 1.0 1.0 1.0 1.0 5.4 4.0 2.0 5.1

Vapor Fraction 0.0 0.0 0.0 0.0 0.0 0.0 1.0 1.0 0.0 0.0 0.0 0.0

Ethanol kg/hr 13 177 190 190 15 1 1 0

Water kg/hr 33,248 329,030 384,784 384,784 356,069 27,158 751 9 4,681 147,140 147,140 17,679

Glucose (SS) kg/hr 0 502 502 502 42 3 0 0

Xylose (SS) kg/hr 0 1,022 1,022 1,022 85 7 0 0

Other sugars (SS) kg/hr 0 2,094 2,094 2,094 175 13 00

Sugar Oligomers (SS) kg/hr 0 1,552 1,552 1,552 130 10 00

Organic Soluble Solids (SS) kg/hr 0 15,808 15,808 15,808 2,387 182 0 0

Inorganic Soluble Solids (SS) kg/hr 2,513 2,513 2,513 110 8 0 0

Ammonia kg/hr 0000 633 48 27 0

Acetic Acid kg/hr 58 58 58 5 0 0 0

Sulfuric Acid kg/hr 0000000

Furfurals kg/hr 329 513 842 842 70 5 1 0

Other Organics kg/hr 0 1,348 1,348 1,348 113 9 0 0

Carbon Dioxide kg/hr 181 14 14,269 3

Methane kg/hr30 5,378 0

O2 kg/hr00010 1,143 30

N2 kg/hr00000550

CO/SOX/NOX/H2S kg/hr000 300 23 285 1

Cellulose (IS) kg/hr 25 25 25 6 19

Xylan (IS) kg/hr88826

Other Struct. Carbohydr. (IS) kg/hr22201

Acetate (IS) kg/hr

Lignin (IS) kg/hr 250 250 250 64 186

Protein (IS) kg/hr 69 69 69 18 51

Cell Mass (IS) kg/hr 19 19 19 280 813

Other Insoluble Solids (IS) kg/hr 92 92 92 23 68

Carbon Mole Flow kmol/hr 20 801 820 820 115 56 660 0

COD kg/hr 636 25,900 26,536 26,536 4,180 1,892 21,994 1

LHV | 25°C Gcal/hr 14.8 100.8 128.3 134.6 187.8 8.9 -65.5 0.0 2.4 87.2 80.8 10.2

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QH602 6.339 WB606 25 0.021

QT606 -1.734 WP602 4 0.004

WT606 200 0.172

126
COMPONENT UNITS 611 612 613 614 616 620 621 622 623 624 625 626 627 628 629 630 632 640

Total Flow kg/hr 360,712 28,626 91,727 120,354 110,595 2,679,447 2,201,450 221,417 9,758 386,274 2,293,177 376,324 9,929 305,040 81,234 223,602 4,504 21

Insoluble Solids % 0.1% 4.0% 1.0% 1.7% 0.1% 0.8% 1.0% 0.0% 20.0% 0.0% 1.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.8% 0.8% 1.3% 1.1% 1.2% 1.3% 1.3% 0.0% 0.9% 1.3% 1.3% 0.0% 49.5% 0.0% 6.0% 0.0% 50.0% 0.0%

Temperature °C 35 35 25 29 29 25 25 25 29 25 25 27 100 25 25 25 20 86

Pressure atm 1.0 1.0 2.0 2.0 1.0 1.0 2.0 1.0 2.0 4.4 1.0 1.0 1.0 1.0 1.0 1.0 2.0 1.0

Vapor Fraction 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.0 0.0 1.0

Ethanol kg/hr 15 1 0 1 1 1 1 0 0 01000

Water kg/hr 356,069 27,158 89,647 116,805 109,098 2,622,465 2,151,522 4,350 7,708 381,300 2,241,169 376,324 4,967 305,040 76,260 4,379 2,252 9

Glucose (SS) kg/hr 42 3 0 3 3 2 2 0 0 02000

Xylose (SS) kg/hr 85 7 0 7 6 4 3 0 0 13110

Other sugars (SS) kg/hr 175 13 0 14 13 8 6 0 1 17110

Sugar Oligomers (SS) kg/hr 130 10 0 10 9 7 6 0 1 16110

Organic Soluble Solids (SS) kg/hr 2,387 182 19 201 187 545 447 0 13 79 466 79 790

Inorganic Soluble Solids (SS) kg/hr 110 8 1,135 1,143 1,068 33,206 27,243 0 75 4,828 28,378 4,828 4,828 2,252 0

Ammonia kg/hr 633 48 1 49 46 19 16 133 16330

Acetic Acid kg/hr 50000000000000

Sulfuric Acid kg/hr 00000000000000

Furfurals kg/hr 70 5 0 5 5 3 3 0 0 03000

Other Organics kg/hr 113 9 0 9 8 8 6 0 1 17110

Carbon Dioxide kg/hr 181 14 2 15 14 44 36 3,857 1 6 38363

Methane kg/hr 30000003000000

O2 kg/hr 10111 19 16 44,765 0 3 170 3 51,0613

N2 kg/hr 00111 37 31 168,157 0 5 320 5 168,1625

CO/SOX/NOX/H2S kg/hr 300 23 10 33 31 304 249 284 2 44 259 44 441

Cellulose (IS) kg/hr 6 19 8 26 1 194 186 25 194

Xylan (IS) kg/hr 26390 65 638 65

Other Struct. Carbohydr. (IS) kg/hr 01120 15 142 15

Acetate (IS) kg/hr

Lignin (IS) kg/hr 64 186 77 263 13 1,925 1,848 250 1,925

Protein (IS) kg/hr 18 51 4 55 3 90 86 52 90

Cell Mass (IS) kg/hr 280 813 791 1,604 80 19,778 18,987 1,523 19,778

Other Insoluble Solids (IS) kg/hr 23 68 28 96 5 707 679 92 707

Carbon Mole Flow kmol/hr 115 56 40 95 12 995 952 88 83 4 991440

COD kg/hr 4,180 1,892 1,347 3,239 448 33,869 32,322 543 2,791 200 33,669 199 2001

LHV | 25°C Gcal/hr 187.8 8.9 48.9 57.6 62.8 1451.4 1174.7 -1.1 -5.2 227.5 1223.7 218.1 8.2 177.3 50.2 0.0 3.3 0.0

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QT608 25.640 WB608 4299 3.696

WC6141 0.000

WM640 2171 1.867

WP609 76 0.065

WP610 301 0.259

WP611 5 0.004

WP612 4 0.004

WP616 47 0.040

WP632 0 0.000

WS611 235 0.202

127
COMPONENT UNITS 309 311 404 443 515 701 703 710 713 735

Total Flow kg/hr 211 948 8 157 21,808 465 22,273 1,981 8,021 1,323

Insoluble Solids % 25.0% 25.0% 25.0% 25.0% 0.0% 0.0% 0.0% 0.0% 0.0% 25.0%

Soluble Solids % 24.8% 24.7% 24.7% 24.7% 0.0% 0.0% 0.0% 0.0% 0.0% 24.7%

Temperature °C 20 20 20 20 38 21 20 21 20 20

Pressure atm 5.4 5.4 1.0 1.0 1.0 6.9 1.0 5.4 1.0 1.0

Vapor Fraction 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

Ethanol kg/hr 21,673 21,673

Water kg/hr 106 476 4 79 108 108 139 8,021 665

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr 52 235 2 39 327

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr00

Acetic Acid kg/hr

Sulfuric Acid kg/hr 1,842

Furfurals kg/hr00

Other Organics kg/hr 465 465

Carbon Dioxide kg/hr 27 27

Methane kg/hr

O2 kg/hr00

N2 kg/hr00

CO/SOX/NOX/H2S kg/hr00

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr 53 237 2 39 331

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr 4 18 0 3 941 32 974 25

COD kg/hr 132 592 5 98 45,162 1,633 46,795 826

LHV | 25°C Gcal/hr -0.4 -1.7 0.0 -0.3 -138.9 -4.9 -143.6 1.2 4.7 -2.4

Work Stream kW Gcal/hr

WP7015 0.004

WP7031 0.000

WP7040 0.000

WP7100 0.000

WP7200 0.000

WT7202 0.002

128
COMPONENT UNITS 273 310 312 406 441 717 755

Total Flow kg/hr 1,051 26 116 6 109 1,166 142

Insoluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 100.0% 100.0% 0.0% 0.0% 0.0% 100.0%

Temperature °C 20 20 20 28 28 20 20

Pressure atm 17.0 1.0 1.0 9.0 9.0 17.0 1.0

Vapor Fraction 0.0 0.0 0.0 1.0 1.0 0.0 0.0

Ethanol kg/hr

Water kg/hr

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr

Inorganic Soluble Solids (SS) kg/hr 26 116 142

Ammonia kg/hr 1,051 6 109 1,166

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr

COD kg/hr

LHV | 25°C Gcal/hr -4.4 0.0 0.0 0.0 -0.5 -4.8 0.0

Work Stream kW Gcal/hr

WC755 0 0.000

WP7600 0.000

WT760 3 0.002

129
COMPONENT UNITS 560 615 802 803 804 806 807 809 810 850 851

Total Flow kg/hr 12,044 21,860 318,399 318,399 284,495 364,696 364,696 5,725 363,445 3,579 895

Insoluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 1.2% 1.2% 100.0% 0.0% 0.0% 100.0%

Soluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 55 38 29 192 25 278 149 0 145 33 25

Pressure atm 6.9 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 5.0 1.0

Vapor Fraction 1.0 1.0 1.0 1.0 1.0 1.0 1.0 0.0 1.0 0.0 0.0

Ethanol kg/hr 1122

Water kg/hr 175 751 6,498 6,498 5,572 50,698 50,698 54,451 3,579

Glucose (SS) kg/hr 0000

Xylose (SS) kg/hr 0000

Other sugars (SS) kg/hr 0000

Sugar Oligomers (SS) kg/hr 0000

Organic Soluble Solids (SS) kg/hr 0000

Inorganic Soluble Solids (SS) kg/hr000 80 80 80

Ammonia kg/hr 0 27 27 27 30 30 30

Acetic Acid kg/hr000

Sulfuric Acid kg/hr000000

Furfurals kg/hr 1122

Other Organics kg/hr 0000

Carbon Dioxide kg/hr 14,269 14,269 14,269 73,396 73,396 73,396

Methane kg/hr 5,378 5,378 5,378

O2 kg/hr 2,764 1,143 68,873 68,873 64,966 11,479 11,479 11,324

N2 kg/hr 9,103 5 223,065 223,065 213,957 223,985 223,985 223,985

CO/SOX/NOX/H2S kg/hr 285 285 285 797 797 178

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr 4,230 4,230 5,725 895

Carbon Mole Flow kmol/hr 0 660 660 660 1,680 1,680 10 1,670

COD kg/hr 4 21,994 21,997 21,997 530 530 326 49

LHV | 25°C Gcal/hr -0.1 -65.5 -65.7 -78.8 0.0 -26.4 -13.3 0.1 -11.6 2.1 0.3

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QH801 13.112 WCOMBFAN 142 0.122

WM804 596 0.512

130
COMPONENT UNITS 215 216 571 623 801 803 806 812 812B 813A 813B 815 815A 815E 816 816A 821 823 840

Total Flow kg/hr 3,490 24,534 36,538 9,758 46,297 318,399 364,696 234,784 30,132 242,044 242,045 75,071 104 101,453 101,453 101,453 4,681 30,132 0

Insoluble Solids % 0.0% 0.0% 62.3% 20.0% 53.4% 0.0% 1.2% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 0.0% 2.5% 0.9% 2.2% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 266 266 55 29 47 192 278 454 266 114 176 233 233 46 46 46 100 192 25

Pressure atm 13.0 13.0 6.9 2.0 1.0 1.0 1.0 59.9 13.0 62.2 62.2 9.5 9.5 0.1 0.1 5.0 5.4 13.0 1.0

Vapor Fraction 1.0 1.0 0.0 0.0 0.0 1.0 1.0 1.0 1.0 0.0 0.0 1.0 1.0 0.9 0.0 0.0 0.0 0.0 1.0

Ethanol kg/hr 6 0 6 2

Water kg/hr 3,490 24,534 12,797 7,708 20,505 6,498 50,698 234,784 30,132 242,044 242,045 75,071 104 101,453 101,453 101,453 4,681 30,132

Glucose (SS) kg/hr 19 0 20 0

Xylose (SS) kg/hr 40 0 40 0

Other sugars (SS) kg/hr 81 1 82 0

Sugar Oligomers (SS) kg/hr 60 1 61 0

Organic Soluble Solids (SS) kg/hr 612 13 625 0

Inorganic Soluble Solids (SS) kg/hr 97 75 173 0 80

Ammonia kg/hr 0 3 3 27 30

Acetic Acid kg/hr 0 0 0

Sulfuric Acid kg/hr 0 0 0 0

Furfurals kg/hr 19 0 20 2

Other Organics kg/hr 52 1 53 0

Carbon Dioxide kg/hr 1 1 14,269 73,396

Methane kg/hr 0 0 5,378 0

O2 kg/hr 1 0 1 68,873 11,479

N2 kg/hr 1 0 1 223,065 223,985

CO/SOX/NOX/H2S kg/hr 2 2 285 797

Cellulose (IS) kg/hr 1,230 25 1,255

Xylan (IS) kg/hr 415 8 423

Other Struct. Carbohydr. (IS) kg/hr 94 2 96

Acetate (IS) kg/hr

Lignin (IS) kg/hr 12,226 250 12,475

Protein (IS) kg/hr 3,376 52 3,428

Cell Mass (IS) kg/hr 925 1,523 2,448

Other Insoluble Solids (IS) kg/hr 4,489 92 4,581 4,230

Carbon Mole Flow kmol/hr 937 83 1,021 660 1,680 0

COD kg/hr 31,489 2,791 34,279 21,997 530 0

LHV | 25°C Gcal/hr -0.3 -2.5 -97.3 -5.2 -102.4 -78.8 -26.4 -42.9 -3.0 119.4 103.9 -6.4 0.0 6.3 56.9 56.9 2.4 12.4 0.0

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QBOILER 144.629 WKNET -12797 -11.003

QH811 15.467 WP811 20 0.017

QM811 50.717 WPLANT -28527 -24.529

WTOTAL -41324 -35.532

131
COMPONENT UNITS 218 250 251 811 811A 811B 811H 812E 813 813A 815A 816A 821 821V 823 896

Total Flow kg/hr 19,782 15,932 3,850 35,284 206,656 241,940 35,284 0 242,044 242,044 104 101,453 4,681 2,580 30,132 75,071

Insoluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 130 100 100 33 114 111 95 0 113 114 233 46 100 100 192 171

Pressure atm 2.6 1.0 1.0 5.4 5.2 4.4 5.4 1.5 1.5 62.2 9.5 5.0 5.4 1.0 13.0 8.0

Vapor Fraction 1.0 1.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.0 0.0 0.0 1.0 0.0 0.0

Ethanol kg/hr 880

Water kg/hr 19,533 15,702 3,831 35,284 206,656 241,940 35,284 242,044 242,044 104 101,453 4,681 2,580 30,132 75,071

Glucose (SS) kg/hr 000

Xylose (SS) kg/hr 000

Other sugars (SS) kg/hr 000

Sugar Oligomers (SS) kg/hr 000

Organic Soluble Solids (SS) kg/hr 000

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr 000

Acetic Acid kg/hr 36 25 11

Sulfuric Acid kg/hr 000

Furfurals kg/hr 205 197 8

Other Organics kg/hr 000

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr 12 11 1

COD kg/hr 396 372 25

LHV | 25°C Gcal/hr -2.9 -1.7 1.9 20.2 101.8 119.9 18.1 0.0 119.5 119.4 0.0 56.9 2.4 -0.1 12.4 32.6

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QH812 -2.185 WP804 8 0.007

WP824 9 0.008

WP826 559 0.481

WP828 2 0.002

132
Work Stream kW Gcal/hr

WM830 10 0.009

WM832 10 0.009

WM834 10 0.009

WP830 1 0.001

133
COMPONENT UNITS 940 941 944 945 948 949 950 951

Total Flow kg/hr 11,923,904 155,041 17,678 11,923,904 6,549,403 6,687,697 1,372,608 1,372,608

Insoluble Solids% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 28 33 28 37 25 37 4 15

Pressure atm 5.0 5.0 5.1 5.0 1.0 1.0 4.1 4.1

Vapor Fraction 0.0 0.0 0.0 0.0 1.0 1.0 0.0 0.0

Ethanol kg/hr

Water kg/hr 11,923,687 155,041 17,678 11,923,687 128,274 266,568 1,372,608 1,372,608

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr 80 0 80 1,495,590 1,495,590

N2 kg/hr 138 0 138 4,925,539 4,925,539

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr

COD kg/hr

LHV | 25°C Gcal/hr 6898.7 89.0 10.2 6801.3 0.0 -20.3 824.1 810.6

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QCHWOP 13.440 WM902 1021 0.878

QTTOTAL 97.401 WM904 186 0.160

WM908 2489 2.140

WP902 1598 1.374

134
COMPONENT UNITS 211 274 400 524 626 811 850 904 905 906 941

Total Flow kg/hr 140,850 150,310 11,419 26,836 376,324 35,284 3,579 147,140 523,463 145 155,041

Insoluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0% 0.0%

Temperature °C 33 33 33 33 27 33 33 59 33 20 33

Pressure atm 5.0 5.0 5.0 5.0 1.0 5.0 5.0 2.0 5.0 1.0 5.0

Vapor Fraction 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

Ethanol kg/hr

Water kg/hr 140,850 150,310 11,419 26,836 376,324 35,284 3,579 147,140 523,463 145 155,041

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr

COD kg/hr

LHV | 25°C Gcal/hr 80.8 86.3 6.6 15.4 218.1 20.3 2.1 80.8 300.5 0.1 89.0

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QT914 1.615 WP914 72 0.062

135
COMPONENT UNITS 906 909 914 916

Total Flow kg/hr 145 145 145 145

Insoluble Solids % 0.0% 0.0% 0.0% 0.0%

Soluble Solids % 0.0% 0.0% 0.0% 0.0%

Temperature °C 20 121 20 20

Pressure atm 1.0 1.0 1.0 1.0

Vapor Fraction 0.0 1.0 0.0 0.0

Ethanol kg/hr

Water kg/hr 145 145 145 145

Glucose (SS) kg/hr

Xylose (SS) kg/hr

Other sugars (SS) kg/hr

Sugar Oligomers (SS) kg/hr

Organic Soluble Solids (SS) kg/hr

Inorganic Soluble Solids (SS) kg/hr

Ammonia kg/hr

Acetic Acid kg/hr

Sulfuric Acid kg/hr

Furfurals kg/hr

Other Organics kg/hr

Carbon Dioxide kg/hr

Methane kg/hr

O2 kg/hr

N2 kg/hr

CO/SOX/NOX/H2S kg/hr

Cellulose (IS) kg/hr

Xylan (IS) kg/hr

Other Struct. Carbohydr. (IS) kg/hr

Acetate (IS) kg/hr

Lignin (IS) kg/hr

Protein (IS) kg/hr

Cell Mass (IS) kg/hr

Other Insoluble Solids (IS) kg/hr

Carbon Mole Flow kmol/hr

COD kg/hr

LHV | 25°C Gcal/hr 0.1 0.0 0.1 0.1

Heat Stream Gcal/hr Work Stream kW Gcal/hr

QH910 -0.091 WM910 250 0.215

136
F1147-E(10/2008) 

REPORT DOCUMENTATION PAGE Form Approved

OMB No. 0704-0188

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1. REPORT DATE (DD-MM-YYYY)

May 2011

2. REPORT TYPE

Technical Report

3. DATES COVERED (From - To)

 

4. TITLE AND SUBTITLE

Process Design and Economics for Biochemical Conversion of 

Lignocellulosic Biomass to Ethanol: Dilute-Acid Pretreatment and 

Enzymatic Hydrolysis of Corn Stover

5a. CONTRACT NUMBER

DE-AC36-08GO28308

5b. GRANT NUMBER

5c. PROGRAM ELEMENT NUMBER

6. AUTHOR(S)

D. Humbird, R. Davis, L. Tao, C. Kinchin, D. Hsu, A. Aden, 

P. Schoen, J. Lukas, B. Olthof, M. Worley, D. Sexton, D. Dudgeon

5d. PROJECT NUMBER

NREL/TP-5100-47764

5e. TASK NUMBER

BB07.2410

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-5100-47764

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NREL

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13. SUPPLEMENTARY NOTES

14. ABSTRACT (Maximum 200 Words)

This report describes one potential biochemical ethanol conversion process, conceptually based upon core 

conversion and process integration research at NREL. The overarching process design converts corn stover to 

ethanol by dilute-acid pretreatment, enzymatic saccharification, and co-fermentation. Building on design reports 

published in 2002 and 1999, NREL, together with the subcontractor Harris Group Inc., performed a complete review 

of the process design and economic model for the biomass-to-ethanol process. This update reflects NREL’s current 

vision of the biochemical ethanol process and includes the latest research in the conversion areas (pretreatment, 

conditioning, saccharification, and fermentation), optimizations in product recovery, and our latest understanding of 

the ethanol plant’s back end (wastewater and utilities). The conceptual design presented here reports ethanol 

production economics as determined by 2012 conversion targets and “nth-plant” project costs and financing. For the 

biorefinery described here, processing 2,205 dry ton/day at 76% theoretical ethanol yield (79 gal/dry ton), the ethanol 

selling price is $2.15/gal in 2007$.

15. SUBJECT TERMS

cellulosic ethanol; production economics; process simulation

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OF ABSTRACT

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