Technical Report

NREL/TP-6A2-46588

June 2010

Techno-Economic Analysis of 

Biochemical Scenarios for 

Production of Cellulosic Ethanol

F. Kabir Kazi, J. Fortman, and R. Anex

Iowa State University

G. Kothandaraman

ConocoPhillips Company

D. Hsu, A. Aden, and A. Dutta

National Renewable Energy Laboratory
National Renewable Energy Laboratory

1617 Cole Boulevard, Golden, Colorado 80401-3393

303-275-3000 • www.nrel.gov

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

Office of Energy Efficiency and Renewable Energy

Operated by the Alliance for Sustainable Energy, LLC

Contract No. DE-AC36-08-GO28308

Technical Report

NREL/TP-6A2-46588

June 2010

Techno-Economic Analysis of 

Biochemical Scenarios for 

Production of Cellulosic Ethanol

F. Kabir Kazi, J. Fortman, and R. Anex

Iowa State University

G. Kothandaraman

ConocoPhillips Company

D. Hsu, A. Aden, and A. Dutta

National Renewable Energy Laboratory

Prepared under Task No. BB07.7510
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iii

Foreword

The purpose of this techno-economic analysis is to compare a set of biofuel conversion 

technologies selected for their promise and near-term technical viability. Every effort has been 

made to make this comparison on an equivalent basis using common assumptions. The process 

design and parameter value choices underlying this analysis are based on public domain 

literature only. For these reasons, the results are not indicative of potential performance. Rather 

they are meant to represent the most likely performance given the current state of public 

knowledge.
iv

List of Acronyms

AFEX ammonia fiber explosion (or expansion)

ASPEN Advanced Simulator For Process Engineering (software)

BTU British thermal unit

CAFI Consortium for Applied Fundamentals and Innovation

COD chemical oxygen demand 

DB declining balance (depreciation)

EtOH ethanol

EVD experimentally validated data

FCI fixed capital investment

FPU filter paper units

GGE gallons of gasoline equivalent

IRS Internal Revenue Service

MACRS (IRS) modified accelerated cost recovery system

PV product value 

MM million

MT metric ton

NREL National Renewable Energy Laboratory

SSCF simultaneous saccharification and cofermentation 

SSF simultaneous saccharification and fermentation 

ST short ton

TD&IC total direct and indirect costs

TCI total capital investment
v

Executive Summary

A techno-economic analysis on the production of cellulosic ethanol by fermentation was

conducted to understand the viability of liquid biofuel production processes within the next 5-8 

years. Initially, 35 technologies were reviewed and a matrix was prepared considering 

economics, technological soundness and maturity, environmental aspects, process performance, 

and technical and economic risks. Then, a two-step down selection was performed to choose 

scenarios to be evaluated in a more detailed economic analysis. In the first screening, the

lignocellulosic ethanol process was selected because it is well studied and portions of the process 

have been tested at pilot scales. In the second screening, seven scenarios of process variations 

were selected: four variations involved pretreatment (dilute acid, two-stage dilute acid, hot water,

and ammonia fiber explosion) and three variations involved downstream processes

(pervaporation, separate 5-carbon and 6-carbon sugar fermentation, and on-site enzyme 

production). Each of these scenarios was examined in detail. Given the time needed for design, 

construction, and startup of large process plants, plants operating in the 5-8 year timeframe 

would likely need to be based on recent experimental data. For this work, process designs were 

constrained to public data published in 2007 or earlier, without projecting for future process 

improvements. Economic analysis was performed for an “nth plant” (mature technology) to 

obtain total investment and product value (PV) (defined as value of the product needed for a net 

present value of zero with a 10% internal rate of return). The final selection among the scenarios 

was performed primarily based on the PV. Sensitivity analysis was performed on PV to assess

the impact of variations in process and economic parameters. Results show that the modeled 

dilute acid pretreatment process without any downstream process variation had the lowest PV of 

$3.40/gal of ethanol (which is $5.15/gallon of gasoline equivalent, GGE) in 2007 dollars. 

Sensitivity analysis shows that PV is most sensitive to feedstock and enzyme costs.

The cellulosic ethanol process is a new technology, for which a pioneer plant is expected to be 

significantly more expensive than the nth plant. To assess the impact of technology maturity on 

pioneer plant cost, a cost growth analysis was performed following a method documented in a

RAND Corporation report. This methodology attempts to incorporate added expenses and startup time for a new process. There is some subjectivity in choosing the parameters for the pioneer 

plant analysis, so a range of parameters was used to estimate pioneer plant costs for three 

scenarios: optimistic, most probable, and pessimistic. The PV obtained from cost-growth 

analysis is substantially larger for a pioneer plant, increasing from $3.40/gal (which is 

$5.15/GGE), before including added expenses, to $5.01/gal ($7.59/GGE), $5.76/gal

($8.72/GGE), and $7.08/gal ($10.71/GGE) for the optimistic, most probable, and pessimistic 

scenarios, respectively.

The PV obtained from the 2007 published data is much higher than the market gasoline price. 

Also, published technological data may not be adequate to accurately project a competitive PV

from a commercial plant of 2000 MT/day capacity. However, this analysis identifies some of the

more cost intensive operations and areas. The current process can only reach approximately

4.5% ethanol in the fermentor beer, which is a third of what grain ethanol plants are achieving

today. The analysis also assumed an enzyme price as $0.69/gal of ethanol produced. Based on 

this analysis, we believe that high-performance enzymes at a cheaper price are required and that 

more research is needed to achieve higher ethanol concentration in the fermentor for this process 

to compete in the current energy market.
vi

Table of Contents

Foreword..................................................................................................................................................... iii

List of Acronyms........................................................................................................................................ iv

Executive Summary.................................................................................................................................... v

Table of Contents....................................................................................................................................... vi

List of Figures ........................................................................................................................................... vii

List of Tables ............................................................................................................................................. vii

Methodology................................................................................................................................................ 1

Summary of the Down-Selection Process..................................................................................1

Project Assumptions..................................................................................................................3

Process Description and Flow Diagram.....................................................................................4

Process Variations......................................................................................................................5

Dilute Acid Pretreatment .....................................................................................................5

Two-Stage Dilute Acid Pretreatment...................................................................................6

Hot Water Pretreatment .......................................................................................................8

AFEX Pretreatment..............................................................................................................8

Separate C5 and C6 Fermentation .......................................................................................9

Ethanol Separation Using Pervaporation...........................................................................10

On-Site Enzyme Production ..............................................................................................12

Methodology for Economic Analysis......................................................................................13

Methodology for Discounted Cash Flow Analysis for nth Plant........................................13

Pioneer Plant Analysis.......................................................................................................14

Results and Discussion............................................................................................................................ 18

nth Plant Cost Analysis.............................................................................................................18

Sensitivity Analysis .................................................................................................................21

Pioneer Plant Analysis Results................................................................................................24

Comparison with Previous Studies......................................................................................................... 25

Conclusions............................................................................................................................................... 27

References................................................................................................................................................. 28

Appendices................................................................................................................................................ 31

Appendix A - Assumptions for Techno-economic Studies of Biochemical Conversion 

Processes............................................................................................................................31

Plant Size, Location, and Construction..............................................................................31

Units...................................................................................................................................31

Feedstock and Enzymes.....................................................................................................31

Material and Energy Balance.............................................................................................32

Equipment Design, Material of Construction, and Costing...............................................33

Chemical Costing...............................................................................................................34

Operating Cost ...................................................................................................................34

Wastewater Treatment Plant..............................................................................................34

Greenhouse Emissions and Control...................................................................................35

Cost Analysis.....................................................................................................................35

Appendix B - Sensitivity Parameters and Values....................................................................37

Appendix C - Sensitivity Results.............................................................................................39

Appendix D - Cost Analysis Result Summary ........................................................................43

Appendix E - Equipment List and Costs, Installation Factors, and Installed Equipment Costs 

for Dilute Acid Pretreatment Processes.............................................................................51
vii

Appendix F - Process Operating Summaries...........................................................................62

Appendix G - General Process Description.............................................................................78

Appendix H - Cost By Area Of Process Scenarios..................................................................81

Appendix I - Down Selection Matrix.......................................................................................85

Appendix J - Analysis: Cost Growth Variables And Results..................................................90

List of Figures

Figure 1. Overall process block diagram of a typical cellulosic ethanol process plant (based on 

NREL’s 2002 design report and modified to 2007 EVD) .................................................. 5

Figure 2. Process flow diagram for dilute acid pretreatment.......................................................... 6

Figure 3. Process flow diagram for two-stage dilute acid pretreatment/hydrolysis and 

fermentation ........................................................................................................................ 7

Figure 4. Process flow diagram for base case pretreatment, enzymatic hydrolysis, and 

fermentation ........................................................................................................................ 7

Figure 5. Hot water pretreatment process flow diagram................................................................. 8

Figure 6. AFEX process flow diagram ........................................................................................... 9

Figure 7. Separate C5 and C6 fermentation configuration ........................................................... 10

Figure 8. Base case distillation configuration............................................................................... 11

Figure 9. Pervaporation separation process.................................................................................. 12

Figure 10. On-site enzyme production process flow diagram...................................................... 13

Figure 11. Impact of pretreatment parameters on PV................................................................... 21

Figure 12. Impact of saccharification parameters on PV.............................................................. 22

Figure 13. Impact of overall process/economic parameters on PV (dilute acid pretreatment)..... 23

Figure 14. Ethanol cost estimations from previous techno-economic studies.............................. 25

Figure G-1. Wastewater treatment section (Area 600)................................................................. 79

Figure G-2. Steam and power generation section (Area 800) ...................................................... 80

List of Tables

Table 1. Plant Performance and Cost Growth Variables for Dilute Acid Pretreatment 

Processes........................................................................................................................... 16

Table 2. Product Value for Various Pretreatment and Downstream Process Variations.............. 18

Table 3. Comparison of Dilute Acid Pretreatment Results from Lab- and Pilot-Scale Data ....... 19

Table 4. Costs by Area of the Dilute Acid Pretreatment Process Scenario .................................. 20

Table 5. Pioneer Plant Analysis Results for the Dilute Acid Pretreatment Process Scenario ...... 24

Table A-1. Corn Stover Feedstock Composition.......................................................................... 32

Table B-1. Sensitivity Parameters for Pretreatment and Saccharification (AREA 200).............. 37

Table B-2. Sensitivity Parameters for Overall Process................................................................. 38

Table C-1. Impact of Pretreatment Parameters on PV.................................................................. 39

Table C-2. Impact of Overall Process and Economic Parameters on PV..................................... 42

Table D-1. Cost Analysis Result Summary for Dilute Acid Pretreatment Processes................... 43

Table D-2. Cost Analysis Result Summary for Dilute Acid Pretreatment (High Solids) 

Processes........................................................................................................................... 44
viii

Table D-3. Cost Analysis Result Summary for Hot Water Pretreatment Process........................ 45

Table D-5. Cost Analysis Result Summary for AFEX Pretreatment Processes........................... 47

Table D-6. Cost Analysis Result Summary for On-site Enzyme Production Processes .............. 48

Table D-7. Cost Analysis Result Summary for Pervaporation Purification Processes................. 49

Table D-8. Cost Analysis Result Summary for Separate C5 and C6 Fermentation Processes..... 50

Table E-1. Equipment Lists and Costs for Dilute Acid Pretreatment Processes.......................... 51

Table F-1. Operating Summary for Dilute Acid Pretreatment Processes..................................... 62

Table F-2. Operating Summary for Dilute Acid Pretreatment (Pilot) Processes.......................... 64

Table F-3. Operating Summary for Two-Stage Dilute Acid Pretreatment Processes.................. 66

Table F-4. Operating Summary for Hot Water Pretreatment Processes....................................... 68

Table F-5. Operating Summary for AFEX Pretreatment Processes............................................. 70

Table F-6. Operating Summary for Separate C5 & C6 Fermentation Processes.......................... 72

Table F-7. Operating Summary for Pervaporation Two-Stage Dilute Acid Pretreatment 

Processes........................................................................................................................... 74

Table F-8. Operating Summary for On-site Enzyme Production Processes................................. 76

Table H-1. Costs by Area of the Dilute Acid Pretreatment (High Solids) Scenario .................... 81

Table H-2. Costs by Area of the 2-Stage Dilute Acid Pretreatment Scenario.............................. 81

Table H-3. Costs by Area of the Hot Water Pretreatment Scenario ............................................. 82

Table H-4. Costs by Area of the AFEX Pretreatment Scenario ................................................... 82

Table H-5. Costs by Area of the Separate C-5 and C-6 Fermentation Process Scenario ............. 83

Table H-6. Costs by Area of the Pervaporation Process Scenario................................................ 83

Table H-7. Costs by Area of the On-site Enzyme Production Processes Scenario ...................... 84

Table I-1. Process Down-Selection Matrix................................................................................... 85

Table J-1. Plant Performance and Cost Growth Variables for Dilute Acid Pretreatment (High 

Solids Loading) Processes ................................................................................................ 90

Table J-2. Plant Performance and Cost Growth Variables for Hot water Pretreatment 

Processes........................................................................................................................... 90

Table J-3. Plant Performance and Cost Growth Variables for Two-Stage Dilute Acid 

Pretreatment Processes...................................................................................................... 91

Table J-4. Plant Performance and Cost Growth Variables for AFEX Pretreatment Processes.... 91

Table J-5. Plant Performance and Cost Growth Variables for On-Site Enzyme Production 

Processes........................................................................................................................... 92

Table J-6. PV from Cost Growth Analysis for Various Process Scenarios.................................. 92
1

Methodology

The objective of this study was to select a biochemical conversion process from published 

literature to produce a second generation liquid biofuel from lignocellulosic biomass. The studies 

involved several steps defining the objective and scope of the project. They started with a review 

of various biochemical processes published in journals and reports and then applied two steps of 

screening to make the final selection. The cost analysis is performed assuming nth plant 

technology (mature technology) and then subjected to cost growth analysis for a pioneer plant. 

The following steps explain the basic methodology for performing the techno-economic analysis:

• Search literature for technologies under consideration

• Perform down selection with criteria to narrow to a few scenarios

• Design process models using AspenPlus process simulation using available experimental 

data

• Size and cost equipment using traditional methods such as literature references and

vendor quotations

• Determine project investments and perform discounted cash flow analysis

• Adjust sensitivity parameters and document results

• Perform pioneer plant cost growth and performance analysis.

Summary of the Down-Selection Process

Thirty-five biochemical fuel production technologies were initially selected for consideration

(Table I-1). Each technology option was evaluated based on economic and technical feasibility, 

environmental performance, and uncertainty criteria to determine, at a high level, the 

technologies with the greatest overall promise in the 5-8 year timeframe. For a plant to be 

operating at commercial levels in this timeframe, the plant would likely need to be designed 

based on current data because of the time needed for design, construction, and startup of a plant. 

A matrix of techno-economic studies was developed for “down selecting” the most promising 

fuels and processes. Major process and economic parameters such as feedstock, capital 

expenditure, operating cost, yield, capacity factor, complexity of the process, level of technology 

development, and internal rate of return were included in the matrix. These parameters cover the 

major process aspects that reflect the overall economics of the processes. This evaluation was 

conducted using past models, if they existed. It was also conducted through publicly available 

literature in 2007. If quantitative numbers were available for the matrix, they were used. But 

since the technologies are not mature, many matrix entries were qualitative. 

Both butanol and ethanol processes were initially considered. Butanol has properties such as 

higher energy density and immiscibility with water that may make it a better transportation fuel. 

However, butanol production processes are currently at the lab scale or very early pilot stage of 

development, and published data on butanol-producing organisms indicate low yields relative to 

ethanol production. Therefore, only ethanol technologies were adopted for analysis. 
2

Ethanol-producing processes were categorized by pretreatment method. Substantial research has 

been done on biomass pretreatment for biochemical conversion at the bench scale by the 

Consortium for Applied Fundamentals and Innovation (CAFI) projects [1, 2, 3]. Based on the 

results of these studies, concentrated acid, SO2-steam, lime, and ammonia-recycle-percolation 

pretreatment were rejected as less attractive due to the high costs associated with the processes.

The pretreatment processes selected for further analysis were: 

1. Acid pretreatment (single-stage dilute and two-stage dilute)

2. Ammonia Fiber Explosion (or Expansion) (AFEX)

3. Hot water pretreatment.

Dilute acid is a pretreatment technology that showed promise in the original CAFI study [1, 2] 

and serves as the base technology for the National Renewable Energy Laboratory’s (NREL’s)

cellulosic ethanol design report [4]. AFEX is a pretreatment that does not require as much water 

as other pretreatments. Hot water pretreatment has low capital investment requirements. Twostage dilute acid is a pretreatment that eliminates the need for enzymes for saccharification, one 

major source of uncertainty in the other pretreatments. Single-stage dilute acid, AFEX, and hot 

water pretreatment conversions are based on bench-scale experiments from CAFI [3]. Two-stage 

dilute acid pretreatments are also based on bench-scale experiments [5, 6].

In addition, process variations in combination with pretreatment that may offer attractive system 

performance were also considered. Among those process variations selected for further study 

was the production of enzymes on-site using hydrolyzate as a carbon source. On-site enzyme 

production may be cheaper by precluding the use of stabilizers and other additives that are 

needed when enzymes are purchased.

Another variation considered was parallel fermentation of 5-carbon (C5) and 6-carbon (C6)

sugars. Current cofermentative organisms do not have high ethanol yields and are not highly 

robust to system variations. We therefore chose to include parallel fermentations of C5 and C6 

sugars for comparison with the baseline fermentation.

Another technology that was considered worthy of further analysis was the use of pervaporation 

instead of the beer distillation column to separate ethanol and potentially allow higher titer and 

substrate utilization in the fermentor. Pervaporation also has the advantage of lower steam and 

utility requirements than a distillation column.

The three downstream technologies—parallel fermentation of C5 and C6 sugars, on-site enzyme 

production, and pervaporation—were compared to a base case using dilute acid pretreatment.

All together, seven scenarios were examined:

• Dilute Acid Pretreatment

• Two-Stage Dilute Acid Pretreatment

• Hot Water Pretreatment
3

• AFEX Pretreatment

• On-site Enzyme Production with Dilute Acid Pretreatment

• Ethanol Separation using Pervaporation with Dilute Acid Pretreatment 

• Separate C5 and C6 Fermentation with Dilute Acid Pretreatment.

Project Assumptions

For each down-selected process a common list of assumptions on process operations and 

economic analysis was made. The scope of the work was to determine the product value of 

cellulosic ethanol for a plant operating in the 5-8 year timeframe. This timeframe was chosen 

because of the renewable fuel standard volumes mandated in the Energy Independence and 

Security Act of 2007 [7]. For a plant to be operating at commercial levels in this timeframe, the 

plant would likely need to be designed based on current data as large process plant projects

typically take more than four years for design, construction, and startup [8]. The major 

assumptions are highlighted below, and the complete list is provided in Appendix A.

• The plant capacity is 2000 MT/day.

• The processes use corn stover as feedstock.

• The feedstock contains 25% moisture and the composition is assumed to be the same as 

that obtained in CAFI II feedstock analysis (Table A-1).

• Plant capacity factor is 96% (350 on-stream days/year).

• Feedstock cost is $83/dry MT ($75/dry ST).

• Purchased enzyme cost from off-site source is $507/MT ($460/ST) of broth of 10% 

protein used at a loading of 31.3 mg protein per gram cellulose in the feed. This cost was 

calculated based on the cost of producing enzyme on-site using hydrolyzate. This cost 

comes to $0.69/gallon of ethanol.

• The processes use 2007 lab-scale, experimentally validated data (2007 EVD) and 

equipment prices (indexed).

• Plant depreciation is calculated following the Internal Revenue Service (IRS) modified 

accelerated cost recovery system (MACRS) over 7 years for the main plant and 20 years 

for the cogeneration area.

• Plant life is 20 years.

• The plant is 100% equity financed.

• Contingency factor and working capital are 20% of total direct and indirect costs 

(TD&IC) and 15% of fixed capital investment (FCI), respectively, for nth plant. 

• The plant initiates operation in 5-8 years.

• Adopted units: cost of all purchased chemicals and feedstocks, plant capacity, and yields 

are reported in metric tons (MT). Operating conditions: temperature (in o

C), pressure (in 

bar), and mass flow rates (in MT/day). 
4

Process Description and Flow Diagram

The different scenarios detailed above were generated by modifying the NREL 2002 production 

process [4]. Appendix G contains a general description of the process steps that have not been 

altered from previous NREL studies. Appendix F contains details of operating conditions for the 

major process steps. The basic cellulosic ethanol production process, modified from NREL’s 

2002 design report [4], comprises nine sections as shown in Figure 1 and listed below.

• Feed Handling (Area 100)

• Pretreatment and Detoxification (Area 200)

• Enzymatic Hydrolysis and Fermentation (Area 300)

• On-site Enzyme Production (Area 400)

• Product Recovery (Area 500)

• Wastewater Pretreatment (Area 600)

• Storage (Area 700)

• Burner/Boiler Turbo-Generator (Area 800)

• Utilities (Area 900) sections.

The AspenPlus Process Simulator is used in process modeling. Current technological data are

used in the simulation and described as the 2007 experimentally validated data (2007 EVD). In 

the pretreatment area, data for AFEX and hot water 2007 EVD are obtained from CAFI II

research [1, 2, 3]. Two-stage dilute acid data are based on available literature [5, 6]. Dilute acid 

pretreatment data are taken from both CAFI II and NREL research. For all other areas, 2007 

EVD conversion data are obtained from NREL research [9]. The overall process block diagram 

is shown in Figure 1. Operating conditions of the major unit operations of each area and mass 

flow rates are included in Figure 1. 
5

AREA 100

FEED HANDLING

(T = 20 o

C; P = 1.01 bar)

AREA 200

PRETREATMENT & 

CONDITIONING

(T = 190 o

C; P = 13.17 bar)

AREA 600

WASTEWATER 

TREATMENT

(T=21.1 o

C, P=1.01 bar)

AREA 700

STORAGE

AREA 800

BURNER/BOILER 

TURBOGENERATOR

(T=510 o

C, P=97.5 bar)

AREA 900

UTILITIES

FEEDSTOCK

WASHED & SHREDDED 

STOVER (3085 MT/d)

ACID

(199 MT/d)

LIME

(147 MT/day)

GYPSUM

(332MT/d)

HYDRO-LYZATE

NUTRIENTS

(109 MT/d)

ENZYME (209 MT/d)

BROTH

VENT

VENT 

(445 MT/d)

EtOH PRODUCT

STILL SOLIDS (1137 MT/day)

EVAP SYRUP (1882 MT/day)

EVAPORATOR CONDENSATE (435 MT/d)

STEAM

(1205 MT/d)

WASTE WATER (1085 MT/d)

VAPOR TO ATM. 

(485 MT/d)

STEAM 

ELECTRICITY (5026 MT/d)

BOILER BLOWDOWN

ANAEROBIC CH4 (7 MT/day)

2667 MT/d

(25% Moisture)

WASH WATER

(755 MT/d)

NUTRIENTS (1MT/d)

WATER

To CLARIFIER

AIR OUT

(118 MT/d) AIR IN

(112 MT/d)

(9987 MT/d) AREA 300

SACCHARIFICATION &

CO-FERMENTATION

(T = 32 o

C, P = 1.01 bar)

SCRUBBER BOTTOM

(436 MT/d)

AREA 500

DISTILLATION 

DEHYDRATION

SCRUBBER 

EVAPORATOR 

(T=100 o

C, P=1.89 bar)

(10301 MT/d)

(440 MT/d)

(2361 MT/d)

RECYCLE WATER (4393 MT/d) 

RECYCLE COND (2381 MT/d)

RECYCLE WATER

(1530 MT/d)

(457 MT/day)

(155 MT/d)

DIGESTOR SOLIDS 

(3 MT/day)

TREATED 

WASTEWATER

(2037 MT/d)

GAS TO STACK

(10544 MT/d)

COMBUSTION AIR

(7794 MT/d)

BOILER FEED WATER

(5181 MT/d)

ASH

(297 MT/d)

COOLING TOWER 

BLOWDOWN

(379 MT/d)

AIR (477 MT/d)

RECYCLE WATER

(2037 MT/d)

VAPOR TO 

ATMOSPHERE 

(37937 MT/d)

MAKEUP WATER

(50854 MT/d)

WATER RETURN

(7754 MT/d)

TO PROCESS

AIR OUT (148 MT/d)

AIR IN (142 MT/d)

Figure 1. Overall process block diagram of a typical cellulosic ethanol process plant (based on 

NREL’s 2002 design report and modified to 2007 EVD)

Process Variations

This section includes a description of each of the process scenarios. Results from the seven 

process variations are shown in Table 2 of the Results and Discussion section.

Dilute Acid Pretreatment

Table F-1 includes a summary of the dilute acid operating conditions and model parameters. The 

biomass from Area 100 is fed by a screw feeder to the presteamer, where low pressure steam 

(163°C, 4.46 bar) is added to maintain a temperature of around 100°C. The presteamer allows a 

portion of the pretreatment heat requirement to be met with low pressure steam. The biomass 

then enters the pretreatment reactor, where high pressure steam (268°C, 13.17 bar) is added as

shown in Figure 2. Sulfuric acid, diluted with process water, is added to the reactor at a rate 

necessary to achieve 1.9 wt% of the liquid phase in the reactor. The reactor temperature, 

pressure, and residence time are maintained at 190°C, 11.6 bar, and 2 minutes, respectively. The 

biomass slurry is then flashed to 1.0 bar in the blow-down tank. The solid fraction is separated 

from the slurry in a Pneumapress pressure filter. In order to reduce toxicity to the fermentation 

organisms downstream, a liming step is added to neutralize excess H2SO4 in the hydrolyzate. 
6

Solid lime is added to the overliming tank along with the aqueous fraction from the Pneumapress 

to raise the pH to 10. The reaction of lime and H2SO4 forms gypsum, which is separated from the 

hydrolyzate as solid cake. The residence time in the overliming tank is 1 hour. The slurry is 

pumped to a second tank where additional H2SO4 is added to reduce the pH to 4.5. The residence 

time in the second tank is 4 hours, which allows the gypsum crystals to grow to an adequate size 

for solid/liquid separation. The gypsum is separated from the slurry in a two-step process, with 

the first being a hydrocyclone and the second being a rotary drum filter. The conditioned 

hydrolyzate is mixed with the solid biomass fraction from which it was previously separated in 

the Pneumapress pressure filter, and the resulting slurry is ready for enzymatic hydrolysis. Mass 

balance and operating conditions of major unit operations are shown in the process flow diagram 

(Figure 2).

ACID

RECYCLED 

CONDENSATE

ACID

STORAGE

STEAM

(1205 MT/day)

VENT (to WWT)

(1073 MT/day)

SOLIDS (1868 MT/day)

LIQUID

ACID

(111 MT/day)

GYPSUM

(332 MT/day)

RECYCLE WATER

(1096 MT/day)

TO AREA 300 

(SACCHARIFICATION)

RECYCLE WATER

(3297 MT/day)

CONDITIONED LIQUID

PREHYDROLYSIS 

REACTOR

BLOWDOWN

TANK

S/L 

SEPARATION

OVERLIMING

NEUTRALIZATION

RESLURRY 

TANK

GYPSUM SEPARATION

LIME (147 MT/day)

ACID SOLUTION

(2469 MT/day, 3.59% conc)

LIGNOCELLULOSICS / 

HYDROLYSATE

SCREW 

CONVEYOR

(3085 MT/day; 54% solids)

AIR (112 MT/day)

T=190 o

C, 

P=11.6 bar

T=101 o

C, 

P=1.01 bar

VENT (to A-800)

(129 MT/day)

T=50 o

C, 

P=1.01 bar

T=57 o

C, 

P=1.01 bar

(7097 MT/day)

LIME 

HANDLING

LIME 

HOPPER 

CAR

VENT

SHREDDED & WASHED 

STOVER (from A-100)

(9987 MT/day)

Figure 2. Process flow diagram for dilute acid pretreatment

Two-Stage Dilute Acid Pretreatment

An operating summary of the two-stage dilute acid pretreatment is provided in Table F-2. In the 

two-stage dilute acid pretreatment process, the first stage solubilizes most of the hemicellulose,

just as in the dilute acid pretreatment process. In the second stage, a higher concentration of acid 

is added to hydrolyze the cellulose and remaining hemicellulose (Figure 3). This contrasts with 

the dilute acid pretreatment process, where enzymes are used to hydrolyze the cellulose (Figure 

4). Mass balance and operating conditions of the major unit operations are shown in the flow 

diagrams (Figures 3 and 4). Two-stage dilute acid conversion data were taken from literature for 

softwood and assumed to be similar for corn stover [5, 6].
7

2ND STAGE

ACID

DILUTE 

SULFURIC ACID

(373 MT/d)

ACID-TREATED 

FEED

(1875 MT/d)

FERMENTOR

CORN STEEP LIQUOR

(104 MT/d)

BEER STORAGE 

TANK

TO A500

DISTILLATION

SEED

FERMENTOR

TO SCRUBBER

(247 MT/d)

OVERLIMING

TANK

GYPSUM

(34 MT/d)

T=210°C

P=18.7 bar

T=32°C

P=1.01 bar

7% contamination

(0.7 MT/d)

T=32°C

P=1.01 bar

(10080 MT/d)

WASTEWATER 

TREATMENT

(351 MT/d)

TO SCRUBBER

(29 MT/d)

DILUTION RECYCLE

(8297 MT/d)

INOCULUM

(5 MT/d)

CORN STEEP LIQUOR

DIAMMONIUM 

PHOSPHATE

Figure 3. Process flow diagram for two-stage dilute acid pretreatment/hydrolysis and fermentation

Detoxified 

Hydrolysate

Enzyme

(209 MT/d)

(9987 MT/d)

Saccharification 

Tank

T = 32 oC

P = 1 atm

Ethanol 

Fermentor

T = 32 o

C

P = 1 atm

CSL

(103 MT/d)

To Fermentor

(9176 MT/d)

To Seed Fermentor

(1020 MT/d) DAP

(0.7 MT/d)

CSL

(5 MT/d)

(979 MT/d)

Inoculum

To Scrubber

(46 MT/d)

To Scrubber

(394 MT/d)

To Beer Tank

Beer 

Storage 

Tank

Scrubber Bottoms

(436 MT/d)

To Distillation

(10301 MT/d)

Seed 

Fermentor

Figure 4. Process flow diagram for base case pretreatment, enzymatic hydrolysis, and 

fermentation
8

Hot Water Pretreatment

Table F-3 contains a summary of model parameters for hot water pretreatment. The chopped and 

washed biomass from Area 100 is mixed with recycled hot water from Area 500. The slurry is 

fed to a plug flow pretreatment reactor. The reactor pressure is maintained at 12.7 bar and the 

temperature is held constant at 190°C. The residence time in the pretreatment reactor is 5 

minutes. The slurry is then cooled to 65°C and flashed to 1.0 bar in the flash tank. Ammonia is 

added to the reactor to neutralize acetic acid formed during the pretreatment process. The xylose 

and cellulose pretreatment yields are shown in Table F-3. Mass flow rates and operating 

conditions of reactor and major units are shown in the process flow diagram (Figure 5).

CONDENSATE

VENT

TO AREA 300 

(SACCHARIFICATION)

BLOWDOWN

TANK

T=65 oC

P=12.4 bar

STEAM

LIGNOCELLULOSICS / 

HYDROLYSATE

SCREW 

CONVEYOR

PREHYDROLYSIS REACTOR

(T = oC, P = 12.43 bar)

AMMONIA

AMMONIA 

TANK

CONDENSATE

AMMONIA

(9973 MT/day)

(13083 MT/d)

T=20 o

C, 

P=17.23 bar

AMMONIA 

NEUTRALIZER

T=80 oC, P=1.01 bar

SHREDDED & WASHED 

STOVER (from A-100)

(3078 MT/day; 54.4% solids)

(13051 MT/d)

(32 MT/day)

(0 MT/d)

Figure 5. Hot water pretreatment process flow diagram

AFEX Pretreatment

A summary of operating parameters for AFEX pretreatment can be found in Table F-4. In the 

AFEX pretreatment process, the biomass is treated with liquid anhydrous ammonia under high 

pressure (17.2 bar) and 60°C for about 5 minutes [10]. The pressure is rapidly released, causing 

the fibers to explode and increasing the access of enzymes to cellulose. Most of the ammonia is 

recovered from the blow-down tank. Residual ammonia is recovered from the solids by a flash 

followed by fractionation from other volatiles. Recovered ammonia vapor is then compressed, 

condensed, and recycled back into the AFEX reactor [10]. The pretreated biomass is slurried into 

a holding tank to be sent to Area 300 for enzymatic hydrolysis (Figure 6).
9

Liquid Ammonia

Shredded Stover

AMMONIA 

STORAGE

Steam

Recycle Water

To Area 300 

(Saccharification)

Biomass Slurry

BLOWDOWN

TANK

Liquid Ammonia

Lignocellulosics/

Hydrolysate

SCREW 

CONVEYOR

AMMONIA 

RECYCLE

COMPRESSOR

AFEX 

REACTOR

Compressed Ammonia

FLASH

DRUM

AMMONIA

RECOVERY Ammonia Vapor

Ammonia Vent

3078.55 MT/d

(53.5% Solids)

1610.78 MT/day

713.67 MT/day

897.1 MT/day

(10461.26 MT/day)

4138.17 MT/day

40.2 MT/day

Lignocellulosics/

Hydrolysate

1736.91 MT/day

4851.84 MT/day

(2401.25 MT/day)

1.5 MT/day

121.79 MT/day

(8061 MT/day)

839.80 MT/day

Condensate

3201.84 MT/d

To Recycle 

(Acid/Other 

HOLDING Volatiles/Water)

TANK

Figure 6. AFEX process flow diagram

Separate C5 and C6 Fermentation

Separate solid and liquid processing (C6 and C5 sugars respectively) takes advantage of 

enhanced yields where xylose is fermented separately using Zymomonas mobilis and glucose is 

fermented separately with yeast (Saccharomyces cerevisiae or Saccharomyces pastorianus). This 

avoids the issue of lower xylose to ethanol yields seen in the current cofermenting organisms. 

One disadvantage of separate processing is that more water is needed to dilute the solids stream 

because the best yields are achieved at low solids loading. The additional water increases ethanol 

recovery costs. The process shown in Figure 7 attempts to mitigate the water issue by using the 

product stream from the liquor (xylose) fermentation to dilute the stream prior to 

saccharification. Table F-5 contains a summary of operating parameters for the separate C5 and 

C6 fermentation scenario, based on work by Dutta et al. [11].
10

ACID

RECYCLED 

CONDENSATE

ACID

STORAGE

STEAM

(1217 MT/d)

VENT (to WWT)

(1068 MT/d)

LIQUID

RECYCLE WATER

(3289 MT/d)

PREHYDROLYSIS 

REACTOR

BLOWDOWN

TANK

NH4OH (61.9%)

ACID SOLUTION

(2443 MT/d, 3.63% conc)

LIGNOCELLULOSICS / 

HYDROLYSATE

SCREW 

CONVEYOR

AIR (112 MT/d)

T=190 o

C, 

P=11.6 bar

T=101 o

C, 

P=1.01 bar

VENT (to A-800)

(118 MT/day)

(7083 MT/d)

SHREDDED & WASHED 

STOVER (from A-100)

(6739 MT/d)

(5671 MT/d)

To WWT

(11 MT/d)

SEPAR

ATOR

(319 MT/d)

H2SO4 (100%)

(203 MT/d)

(7402 MT/d)

Seed 

Fermentor

C-5 Sugar 

Fermentor

T = 32 oC

P = 1.01 bar

RECYCLE 

WATER

(1627 MT/d)

To SCRUBBER 

(17 MT/d)

SEPAR

ATOR

(93 MT/d)

CSL

Saccharification/ 

Fermentor

T = 32 oC

P = 1.01 bar

(209 MT/d)

WATER

(3079 MT/d; 54.4% solids)

YEAST

CSL

C-6 Sugar 

Fermentor

T = 32 oC

P = 1.01 bar

SEPAR

ATOR

(16 MT/d)

(114 MT/d)

(11379 MT/d)

To SCRUBBER

(304 MT/d)

To BEER 

STORAGE TANK

(11075 MT/d)

T=32 o

C

P=1.01bar

SEPARATOR

(912 MT/d)

(923 MT/d)

(929 MT/d)

DAP

(0.7 MT/d)

CSL

(5 MT/d)

(8310 MT/d)

(9315 MT/d)

To SCRUBBER 

(134 MT/d)

(9180 MT/d)

(1860 MT/d)

(11249 MT/d)

REACIDIFICATION 

TANK

T=63 o

C, 

P=1.01 bar

NEUTRALIZATION 

TANK

T=30 oC, 

P=1.01 bar

S/L 

SEPARATION

Figure 7. Separate C5 and C6 fermentation configuration

Ethanol Separation Using Pervaporation

Pervaporation refers to separation using a membrane with liquid feed on one side and a lowpressure, gaseous permeate output on the other side. Components in the liquid feed preferentially 

permeate through the membrane and then evaporate into the gaseous phase. Because 

pervaporation does not involve a large heat input, the process could save on costs associated with 

the heat and steam needed for the reboiler of a conventional distillation column. A summary of 

operating parameters for the pervaporation separation scenario is located in Table F-6.

In contrast to the base case distillation configuration (Figure 8), the pervaporation variation 

modifies the distillation section of the dilute acid model by inserting a pervaporation system in 

place of the beer column (Figure 9). The pervaporation system output is calculated from a 

separation factor and a total material flux needed to achieve the same separation as in the beer 

column [12]. The separation factors and material fluxes are based on literature [12]. In the base 

case dilute acid model, the beer column also serves to separate carbon dioxide from the ethanol 

stream. In the pervaporation model, a flash tank is added to separate out CO2 at 110°C and a heat 

exchanger cools the stream to 41°C. The separation of CO2 is assumed to be easier than what is 

modeled here, so neither the capital nor utility costs of the flash tank and heat exchanger are 

included in the economic analysis. The membrane system costs $200/m2 in 1999 dollars, with a 

replacement needed every five years at a cost of $100/m2 [13].
11

(10301 MT/d)

P-378

From Beer Storage

Beer Storage 

Tank

To Scrubber

(14 MT/d)

Rectification 

Column

99.5% EtOH

(457 MT/d)

93% EtOH

(612 MT/d)

Steam

Vapor

1st Effect 

Evaporator 2nd Effect 

Evaporator

2nd Effect 

Evaporator

S/L 

Separation

Solids to Combustor

Vapor Evaporator Condensate

(1426 MT/d)

To Combustor

(1882 MT/d)

(1335 MT/d) (878 MT/d)

(155 MT/d)

Regenerant

Molecular 

Sieve

Rectification Bottom to Pneuma Press

(7467 MT/d)

(3308 MT/d)

Liquor

CIP CIP

CIP

(1137 MT/d)

To A-200 

(Pneuma Press)

(4743 MT/d)

Vapor Entrainment

(6 MT/d)

(1581 MT/d)

Condensate

(1484 MT/d)

Condensate

To A200

(8951 MT/d)

(1484 MT/d) (1435 MT/d)

Condensate

(1435 MT/d)

Figure 8. Base case distillation configuration
12

(10637 MT/d)

P-358

From Beer Storage

Pervaporation

Rectification 

Column

99.3% EtOH

(461 MT/d)

92% EtOH

(617 MT/d)

Steam

Vapor

1st Effect 

Evaporator

2nd Effect 

Evaporator 2nd Effect 

Evaporator

S/L 

Separation

Solids to Combustor

Vapor Evaporator Condensate

(1483 MT/d)

To Combustor

(1898 MT/d)

(1475 MT/d)

(1014 MT/d)

(156 MT/d)

Regenerant

Molecular 

Sieve

Rectification Bottom to Pneuma Press

(7625 MT/d)

(3381 MT/d) (4876 MT/d

Liquor )

CIP CIP CIP

31% EtOH

To A-200 

(Pneuma Press)

(1118 MT/d)

Vapor Entrainment 

(6 MT/d)

Condensate Condensate

(1537 MT/d)

Condensate

(1495 MT/d)

(1537 MT/d)

(9162 MT/d)

To A200

(1625 MT/d)

(1495 MT/d)

Figure 9. Pervaporation separation process

On-Site Enzyme Production

Table F-7 contains a summary of the on-site enzyme production operating parameters. On-site 

enzyme production (Area 400) is modeled as a variation to the baseline dilute acid process model 

(where enzyme is assumed to be purchased from suppliers). This work was based on a design 

report by NREL in 1999 [14]. Cellulase enzyme is a mixture of endoglucanases, exoglucanases,

and β-glucosidase enzymes. In the present case, Trichoderma reesei (a fungal strain) is used for 

on-site cellulase enzyme production. A portion of the conditioned pretreated biomass from Area 

200 is pumped to Area 400, where a fraction of the stream is used for the growth of Trichoderma 

reesei inoculums in seed bioreactors. The remainder is pumped into the jacketed aerobic 

bioreactors where inocula are added from the seed bioreactors to produce enzymes. Corn steep 

liquor and other trace nutrients are also added to the reactors. Ammonia is used to control pH and 

to provide additional fixed nitrogen to the fungus; oxygen is supplied by an air compressor. To 

control foam formation, corn oil is added into the bioreactors. Mass flow rates and process 

conditions of major unit operations are shown in Figure 10.
13

(1306 MT/d)

CSL (1 MT/d)

SEED 

REACTOR

DILUTE DETOXIFIED HYDROLYZATE (118 MT/d)

CSL

NUTRIENTS

ANTIFOAM

MEDIA PREPARATION ANTIFOAM TANK

TANK

CELLULASE 

FERMENTOR

AMMONIA/SO2

(6 MT/d)

VENT

(895 MT/d)

DETOXIFIED HYDROLYZATE 

AND LIGNOCELLULOSIC

VENT

(11396 MT/d)

CSL/NUTRIENTS (19.6 MT/d)

AIR

(873 MT/d)

AIR (11212 MT/d)

T=28 o

C, 

P=1.01 bar

(8.7 MT/d)

T=28 o

C, 

P=1.68 bar

T=20 o

C, 

P=1.01 bar

T=20 oC, 

P=1.01 bar

AIR

COMPRESSOR

(1358 MT/d)

CELLULASE TO 

AREA-300

AIR

(12085 MT/d)

AMMONIA/SO2 (0.3 MT/d)

Figure 10. On-site enzyme production process flow diagram

Methodology for Economic Analysis

The base case cost estimation and analysis is performed assuming the plant is an nth plant design.

This means that a similar plant was previously constructed and operated without unexpected 

delays in startup and capacity loss. The process design is assumed to become the nth plant, 

operating at the reported yields from experimental, bench-scale data. This is a major assumption 

for a process that is still early in its development, and when an nth plant for cellulosic ethanol 

becomes commercial, it may not look like this design. However, the nth plant analysis was 

chosen to provide an analysis parallel to analyses done by different groups, including NREL, the 

CAFI group, and Pacific Northwest National Laboratory.

However, because of the new technological elements of cellulosic ethanol production there are a 

number of engineering design and performance uncertainties. These uncertainties are accounted 

for in the cost growth analysis for a pioneer plant using the risk analysis methodology developed 

by RAND [15, 16]. The details of the cost estimation for the nth plant and the cost growth 

analysis for the pioneer plant are discussed in the following subsections.

Methodology for Discounted Cash Flow Analysis for nth Plant

Mass and energy balances for each of the seven selected process variations are performed by 

AspenPlus Process Simulator. Pinch analysis is performed to optimize energy balances. The 
14

stream flow rates, from AspenPlus simulations, are used to size unit operations. The costs for

most of the equipment are obtained from previous vendor quotes obtained by NREL. Individual 

equipment is scaled and the scaled cost is estimated following exponential correlations as 

described in Appendix A. The scaled cost is then indexed to a year 2007 dollar value using the 

Chemical Engineering Plant Cost Index [17]. Separate installation factors are used for each of 

the unit operations to obtain individual installed equipment costs. The installation factors are 

obtained from the vendors who provided equipment quotes. The cost analysis is performed 

following NREL’s approach [4] and that found in Peters et al. [18], with modified terminologies. 

Total installed cost (TIC) is defined as the sum of total installed equipment cost, warehouse, and 

site development costs.

Engineering and supervision costs, construction expenses, and legal and contractor’s fees are 

assumed to be 32%, 34%, and 23% of purchased equipment cost, respectively. Total direct and 

indirect costs (TD&IC) is the sum of total installed cost and indirect costs. Contingency is 

assumed to be 20% of TD&IC, and working capital is 15% of fixed capital investment (FCI). 

FCI is the sum of TD&IC and contingency, and the total capital investment is the sum of FCI

and working capital.

The manufacturing costs include raw materials costs (such as corn stover), variable costs (such 

as process chemicals, enzyme, nutrients) and fixed operating costs (employee salaries, overhead, 

insurance, and maintenance). The feedstock cost is assumed to be $83/dry MT ($75/dry ST)

(Appendix A). The chemical and nutrients costs are obtained from NREL’s previous quotes from 

suppliers and are indexed to 2007 dollar values following the Inorganic Chemical Index of the 

SRI International Economics Handbook, Economic Environment of the Chemical Industry [19]. 

NREL estimated the required number of employees and their salaries in their 2002 design report

[4]. In the present study, the same number of employees is required for the same plant capacity, 

and the salaries are indexed to 2007 dollar values following the labor index from the Bureau of 

Labor Statistics [20].

In the present analysis, the process and steam generation plants are assumed to depreciate in 7 

and 20 years, respectively, following IRS MACRS, and the plant life is 20 years. The project is 

assumed to be 100% equity financed and internal rate of return is 10%. For the present cost 

analysis, the capital investment is spread over 3 years at a rate of 8%, 60%, and 32% in the first, 

second, and third years, respectively. The product value (PV) of ethanol is calculated by iterating

to reach a net present value of $0 with a 10% internal rate of return.

Pioneer Plant Analysis

The pioneer plant analysis is performed using a method developed by Merrow et al. of RAND 

Corporation [15]. This methodology considers two sources of production cost growth in pioneer 

plants: less than expected plant performance and low capital cost estimation. These sources are 

regressed with two multi-factor ordinary least squares correlations to estimate the unexpected 

reduced plant performance (Equation 1) and capital cost growth (Equation 2) associated with 

pioneer plants.

Equation 1 estimates pioneer plant performance as a percentage of design capacity in the second 

6 months after startup. 
15

Plant Performance = 85.77 – 9.69×NEWSTEPS + 0.33×BALEQS

 – 4.12×WASTE – 17.91×SOLIDS (Eq. 1)

Where, 

NEWSTEPS ≡ The number of steps in the process that have not been proven 

commercially.

BALEQS ≡ The percentage of mass and energy balance equations used in plant design

that are validated with commercial-scale data. The RAND report also mentions 

that some weight is given to rigorous theoretical models.

WASTE ≡ Potential problems that may be associated with waste handling. A 0-5 scale is 

used, with 0 meaning no waste handling issues and 5 meaning significant waste

handling issues.

SOLIDS: The scale used is 0 or 1. If the process handles solids then the value is 1;

otherwise it’s 0.

Equation 2 estimates the capital cost growth, defined as the ratio of estimated to actual costs. 

Cost Growth = 1.12196 – 0.00297×PCTNEW – 0.02125×IMPURITIES 

 – 0.01137×COMPLEXITY + 0.00111×INCLUSIVENESS 

 – C1×PROJECT DEFINITION (Eq. 2)

Where,

PCTNEW ≡ The installed cost of all commercially undemonstrated equipment as

percentage of total installed equipment cost.

IMPURITIES: Represents the potential process issues that may arise due to impurity 

buildup from recycle streams or problems due to equipment corrosion. The value

ranges from 0 to 5, with 0 being given to processes with no impurity buildup or 

corrosion issues.

COMPLEXITY ≡ The number of continuously linked process steps.

INCLUSIVENESS ≡ The percentage of three factors: pre-startup personnel costs, prestartup inventory cost, and land purchase. For example, if two of these factors 

have been rigorously considered, the variable would be given a value of 67%.

C1: C1 is 0.06361 if the design is at pre-development/exploratory or research and 

development stage and 0.04011 if the design is in commercial or pre-commercial 

stage. For the present studies C1 is assumed to be 0.06361. 

PROJECT DEFINITION: Includes commitment of funds to define the plant scope, basic 

plant layout, and process flow conditions. Most major equipment is defined and 

examination of site begins at this point. The amount of work involved here 

depends on how much information is already available from previous project

experience. Often some critical level of engineering (heat and mass balances, 

equipment need, and so forth) and site-specific information (on-site and off-site 

unit configurations, soils and hydrology data, health and safety requirements, and 

environmental requirements) is completed here. A numerical value is assigned to

define the level of engineering completed at the time of estimation, following the 
16

level of completeness: (1) engineering completed, (2) moderate or extensive 

engineering, (3) limited engineering, and (4) screening design stage. Similarly, a 

value for site-specific information is assigned by the following: (1) definitive or 

completed work, (2) preliminary or limited work, (3) assumed or implicit 

analysis, and (4) not used in the cost estimation at all. The range of values given 

to project definition is 2 (for maximum definition) to 8 (for no definition).

Cost growth analysis for all seven process variations was performed. There was some 

subjectivity in choosing the parameters for the pioneer plant analysis, so a range of parameters 

was used to estimate pioneer plant costs for three scenarios: optimistic, most probable, and 

pessimistic. For the dilute acid pretreatment process the selected variable values of Equations 1-2 

are shown in Table 1; the selection justification is discussed below. For all other process 

variations, the variable values are shown in Appendix J. 

Table 1. Plant Performance and Cost Growth Variables for Dilute Acid Pretreatment Processes

Plant Performance (Equation 1) Cost Growth (Equation 2)

Variables

Values

Variables

Values

OptimisticMost 

Probable

PessimisticOptimisticMost 

Probable

Pessimistic

NEWSTEPSa 6 6 6 PCTNEW 61.76 61.76 61.76

BALEQS 50 40 30 IMPURITIES 0 3 5

WASTE 1 2 3 COMPLEXITYb 6 6 6

SOLIDS 1 1 1 INCLUSIVENESS 33 0 0

PROJECT 

DEFINITION 6 6 7

Plant 

Performance (%) 22.10 14.68 7.26 Cost Growth 0.53 0.42 0.32

a New steps/units: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, Beer Column, and 

Combustor. bContinuously linked steps: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, 

Distillation, and Steam/Power Generation.

Justification of Correlation Variable Value Selection for Plant Performance (Equation 1)

For all three scenarios, the operations considered new steps/units are feedstock handling, 

pretreatment, saccharification, cofermentation, beer column, and the fluidized bed combustor

(for converting lignin to heat and power), resulting in a value of 6 for NEWSTEPS.

Some of the steps and units are being used commercially, so the BALEQS variable is assigned a 

value of 50, 40, and 30 for optimistic, most probable, and pessimistic scenarios, respectively.

The wastewater contains a small amount of furfural that may not be degraded by the anaerobic or 

aerobic treatments used in the model, meaning that an additional chemical treatment may be 

necessary. No additional complications with waste are foreseen. The WASTE variable was 

assigned a value of 2 for the most probable case. Values of 1 and 3 were assigned for the 

optimistic and pessimistic cases, respectively.
17

These variable values are used in Equation 1 to calculate the percentage of Plant Performance for 

the three cases (Table 1). 

Justification of Correlation Variable Value Selection for Cost Growth (Equation 2)

The feedstock handling area, pretreatment area, saccharification, cofermentation, beer column,

and fluidized bed combustor are selected as new technologies/units to calculate the parameter 

PCTNEW for all three of the cases. 

Some of the degradation products inhibit the saccharification and fermentation process, and 

buildup of those inhibitors in the process may result in yield loss. For the most probable case, the 

assigned value for the variable IMPURITIES is 3. For the optimistic and pessimistic cases the 

assigned values are 0 and 5, respectively. 

The process design has six continuously linked process steps, which include feedstock handling, 

pretreatment, saccharification, cofermentation, distillation, and steam/power generation. 

Therefore, the value for the variable COMPLEXITY is assigned as 6 for all three cases. 

Some of the initial plant inventory is included in the base case cost estimate, although it is not 

validated in a commercial plant. So, for the optimistic scenario a value of 33% is assigned for the 

variable INCLUSIVENESS. The value of 0% is assigned for both pessimistic and most probable

cases. 

A plant site has not been chosen, so none of the site-specific information has been procured. 

Some level of engineering has been completed. Therefore, a value of 6 is assigned for most 

probable and optimistic cases for the variable PROJECT DEFINITION and the pessimistic case 

is assigned a 5. 

The assigned variable values are used in Equation 2 to calculate the percentage of cost growth as 

shown in Table 1. The base case nth plant TCI is divided by the cost growth, obtained from 

Equation 2 (Table 1), to obtain an estimate for the pioneer plant TCI. The contingency factor is 

increased to 30%, compared with 20% for an nth plant, to account for additional construction 

uncertainties. The plant performance, obtained from Equation 1 (Table 1), is multiplied by the 

first year ethanol sales to account for the reduced production of a pioneer plant. For the 

discounted cash flow analysis, plant performance is increased by 20% per year until design 

capacity is reached. 
18

Results and Discussion 

nth Plant Cost Analysis

Ethanol yield, byproduct credit, total installed equipment cost, total project investment, and 

estimated PV for each of the process variations are shown in Table 2.

Table 2. Product Value for Various Pretreatment and Downstream Process Variations

Process 

Variations

Total 

Capital 

Investment 

($MM)

Total 

Installed 

Equipment 

Cost ($MM)

Ethanol 

Yield 

(Gal/MT)

a

Ethanol 

Production 

(MM 

Gal/Yr)

b

Electricity 

Export 

($MM/Yr)

Product 

Value 

($/Gal)

c

Dilute Acid 

Pretreatment 

(base case)

376 164 76.3

(288.8) 53.4 (202.2) 11.7 3.40(0.90)

Dilute Acid 

Pretreatment 

(high solids)

389 169 72.5

(274.5) 50.8 (192.1) 12.6 3.60(0.95)

Two-Stage Dilute 

Acid Pretreatment 391 173 46.8(177.5) 32.8 (124.2) 16.8 4.38(1.16)

Hot Water 

Pretreatment 361 156 55.8(211.0) 39.0 (147.7) 11.3 4.44(1.21)

AFEX 

Pretreatment 386 167 65.9(249.7) 46.2 (174.8) 16.9 3.69(0.97)

PervaporationDistillation 501 209 76.9

(291.3) 53.9 (203.9) 13.6 3.75(0.99)

Separate C5 and 

C6 Fermentation 386 168 79.3 (300) 55.5 (210.1) 6.5 3.67(0.97)

On-site Enzyme 

Production 434 188 67.7(256.3) 47.4 (179.4) -0.8 3.54(0.94) a Values in parentheses are in liter/MT. b Values in parentheses are in MM liter/year. c Values in parentheses are in 

$/liter.

Each pretreatment process has some variation in yield (47-76 gal/MT) with dilute acid 

pretreatment being the highest. The lowest PV among all process variations is from the dilute 

acid pretreatment scenario, which is $3.40/gal, and those for other pretreatment processes are in 

the range of $3.60/gal to $4.44/gal. The yield from two-stage dilute acid pretreatment is lowest

(46.8 gal/MT) and the process requires higher TCI ($391 million), which drives the PV as high 

as $4.38/gal.

The installed equipment cost and TCI for all four pretreatment scenarios are in the range of 

$156-$173 million and $361-$391 million, respectively, with hot water pretreatment being the 

lowest and two-stage dilute acid pretreatment being the highest. The reason for lower installed 

equipment cost for hot water is because it uses a relatively simple horizontal tubular pretreatment 

reactor. The installed cost of the tubular reactor is $311,000 for hot water pretreatment, whereas 

pretreatment reactors for dilute acid and AFEX pretreatment processes are $22.99 million and

$9.15 million, respectively. The dilute acid pretreatment process requires long retention time for 
19

overliming, which requires large expensive vessels. Although the AFEX reactor cost is lower 

than the dilute acid pretreatment reactor, the AFEX pretreatment process has additional 

expensive unit operations (such as the ammonia compressor) that increase the total installed 

equipment cost to slightly more than that of dilute acid pretreatment processes.

For the comparison of pretreatment technologies, lab-scale experimental data at low solids 

loading (25%) were used in the model [3]. In order to understand how process scale-up may 

impact ethanol production cost, a process model was developed using data from experiments 

conducted at NREL at higher solids loading (40%). Table 3 presents a summary comparison of 

the results from the dilute acid pretreatment models.

Table 3. Comparison of Dilute Acid Pretreatment Results from Lab- and Pilot-Scale Data 

Base Case 

(2007 EVD) High Solids

Solids Loading 25% 40%

PV ($/gal) 3.40 3.60

Ethanol Yield (gal/MT) 76.3 72.5

Installed Equipment Cost (MM$) 164.1 169.4

Fixed Capital Investment (MM$) 326.8 337.8

Total Capital Investment (MM$) 375.9 388.5

Lang Factor 3.44 3.43

The PV of the model using high solids loading is $0.20/gal higher than that of the base case 

model. This is primarily due to the decreased ethanol production caused by lower yields of 

monosaccharides at higher solids loadings. 

The costs of each process area for the dilute acid pretreatment scenario are presented in Table 4;

costs for other scenarios are given in Appendix H. The expensive areas of the dilute acid 

pretreatment scenario are the pretreatment, saccharification and fermentation, distillation and 

solids recovery, and boiler and turbo-generator sections. Among these, the boiler and turbogenerator section is the most expensive area, accounting for 56% of the total installed equipment 

costs. 
20

Table 4. Costs by Area of the Dilute Acid Pretreatment Process Scenario

Cost Areas/Factor

Installed Cost Purchased Equipment 

Cost

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 6.6 6.0 5.5

Pretreatment (Area 200) 36.2 22.1 19.7 18.0

Saccharification and Fermentation (Area 300) 21.8 13.3 17.3 15.8

Distillation and Solids Recovery (Area 500) 26.1 15.9 17.2 15.8

Wastewater Treatment (Area 600) 3.5 2.1 2.6 2.4

Storage (Area 700) 3.2 2.0 2.0 1.8

Boiler/Turbogenerator (Area 800) 56.1 34.2 40.3 36.9

Utilities (Area 900) 6.3 3.8 4.2 3.8

Total Installed Equipment Cost 164.1 100 109.3 100

Fixed Capital Investment (FCI) 326.8

Working Capital (WC) 49.0

Total Capital Investment (TCI) 375.9

Lang Factora

3.44

a The Lang factor is calculated by dividing TCI by the total purchased equipment cost.

The PV from the pervaporation scenario is $3.75/gal, which is higher than the PV for the base 

case process (dilute acid pretreatment process, $3.40/gal). The cost differential comes from the

high capital cost of the pervaporation membrane, which is not well developed commercially. 

Ethanol yield in the dilute acid pretreatment process with cofermentation is 76.3 gal/MT. It was 

thought that separate C5 and C6 sugar fermentation using selective yeast might reduce the PV

because of higher yields (79.3 gal/MT). However, the PV is 8% higher than the base case

scenario because of increased capital costs primarily due to the additional fermentation vessels 

required.

It was also thought that on-site enzyme production might provide economic advantages over 

purchasing enzymes because it eliminates the need to use stabilizing chemicals and to 

concentrate the enzyme broth prior to transportation. The PV from the on-site enzyme production 

process is $3.54/gal, which is $0.14/gal higher than the base case dilute acid pretreatment 

process. The cost of the enzyme is affected by a lower electricity credit than the base case. This 

is because of high electricity consumption by the compressor supplying air to the enzyme 

production bioreactors that leads to lower net excess electricity and lower ethanol yield (67.7

gal/MT). In the on-site enzyme production process, part of the feedstock (hydrolyzate from Area 

200) is diverted to enzyme production, which reduces the overall plant capacity by 6 MMGal of 

ethanol per year. It is assumed that Trichoderma reesei is the enzyme-producing fungal strain. 

The specific activity of enzyme, yield, and productivity are 600 FPU/g protein, 0.33 g protein/g 

cellulose and xylose, and 0.125 g protein/L-hr, respectively. The PV is higher for the on-site case 

than for the off-site case because the off-site case’s enzyme cost was determined after increasing 

the feedstock in order to produce the same amount of ethanol. This increase in scale provided 

some economic benefit. This comparison does not include costs of stabilizing chemicals 
21

associated with purchased enzymes, so the enzyme cost continues to be uncertain. However, the 

comparison of an on-site case helps demonstrate what yields and electricity tradeoffs occur when 

enzyme is produced on-site. 

Sensitivity Analysis 

Process-specific sensitivity analysis of pretreatment and saccharification operations was

performed on all pretreatment process scenarios. This analysis showed the impact of process 

operation parameters including operating temperature, retention time, acid concentrations, and 

yields on PV. The results are shown in Figures 11 and 12 (the detailed parameter values and 

results are tabulated in Table B-1 and Table C-1, respectively). PV is most sensitive to 

pretreatment solid consistency, retention time, and xylan and cellulose conversions. When the 

retention time of the dilute acid pretreatment reactor is increased from 2 to 10 minutes, an 

increase in PV of 16% is observed (Figure 11). When the conversion of xylan to xylose in the 

pretreatment reactor is reduced from 82.5% (2007 EVD) to 33%, the PV increases by 6%. And 

when the solid consistency in the hot water pretreatment reactor is increased from 13% to 20%,

the PV is reduced by 10%. Xylose and cellulose conversions in the pretreatment reactor of the 

two-stage dilute acid scenario showed significant impact on PV. The PV increased by 44% and is 

reduced by 10% when xylan to xylose conversion is reduced from 82.5% to 33% and cellulose to 

glucose conversion is increased from 6.3% to 23%, respectively. The impact of other 

pretreatment parameters on PV is not so significant.

Figure 11. Impact of pretreatment parameters on PV
22

Among the saccharification parameters, cellulose to glucose conversions showed significant 

impact on PV for all scenarios (Figure 12; the detailed parameter values and results are tabulated 

in Table B-1 and Table C-1, respectively). For the dilute acid pretreatment process, the PV

increased by 20% when the cellulose to glucose conversion was reduced from 91% to 67%. The 

impact on PV for the hot water pretreatment scenario was much higher. For the hot water 

scenario, the PV increased by 31% when the cellulose to glucose conversion was reduced from 

90% to 65%. 

Figure 12. Impact of saccharification parameters on PV

A sensitivity analysis for the major economic assumptions was performed on the dilute acid 

pretreatment process to test the robustness of the estimated PV. The selected sensitivity 

parameters were feedstock cost, enzyme loading, enzyme cost (for purchased enzymes), 

contingency factor, installation factor (or corresponding installed equipment cost), and byproduct

(electricity) credit. The feedstock cost and enzyme price have the dominant impact on PV. 

Enzyme loading, contingency factor, and the total installed equipment cost showed moderate 

impact on PV. Results are shown in Figure 13 (the detailed parameter values and results are 

tabulated in Table B-2 and Table C-2, respectively).
23

Product Value, PV ($/GGE)

Feedstock Cost (55 : 83 : 110 $/MT)

Enzyme Cost (256 : 507 : 1460 $/MT of broth)

Enzyme Loading (20 : 31.3 : 33.3 mg/g protein)

Contingency (10 : 20 : 30 % TPI) 

Pretreatment Reactor (23, + 25% MM$)

Installed Equipment Cost (159 : 164 : 194 MM$)

Electricity Price (0.06 : 0.054 : 0.04 $/kWh)

$4.15 $4.65 $5.15 $5.65 $6.15 $6.65 $7.15 $7.65

_

Figure 13. Impact of overall process/economic parameters on PV (dilute acid pretreatment)

When feedstock cost was increased from $83/dry MT (base case scenario) to $110/dry MT, PV

increased by 11%. PV increased by 38% to $4.70/gal when enzyme cost was increased from 

$507/MT broth to $1460/MT (equivalent to $2.00/gal of ethanol produced). When enzyme cost 

was reduced to $256/MT (equivalent to $0.35/gal of ethanol produced), a 10% decrease in PV

was observed. The exact cost of cellulase enzymes for large scale needs is not yet publicly 

available. A recent Novozymes presentation estimated enzyme costs in 2009 to be around $1–

$2/gal of ethanol, and this range is included in the sensitivity analysis [21]. It may be important 

to further study the on-site enzyme production process with emphasis on microbial strain, protein 

yield, specific activity, residence time, oxygen requirement, and overall process area 

optimization. The current study is limited to publicly available data on those important 

parameters. 

The enzyme loading in the saccharification reactor showed some impact on PV. When the 

enzyme loading was reduced from 31.3 mg protein/g cellulose (equivalent to 18.8 FPU/g 

cellulose) to 20 mg protein/g cellulose (equivalent to 12 FPU/g cellulose), the PV was reduced 

by 7%; the PV increased by 1% when the loading was increased from 31.3 to 33.3 mg protein/g 

cellulose (equivalent to 20 FPU/g cellulose).

Sensitivity analysis on total installed equipment cost was also performed. When the installed 

equipment cost was increased from $MM164.1 (corresponding weighted average installed factor 

of 2.58 for the base case scenario) to $MM194 (corresponding installed factor of 3.05 obtained 

from Peters and Timmerhaus [18]), the PV increased by 4%. Contingency factor showed a 

similar impact on PV. The impact of other parameters such as reactor cost and electricity price 

on PV was not significant.
24

Pioneer Plant Analysis Results

Table 5 shows the pioneer-plant PV results for all three cases. The PV for the most-probable case 

is $5.76/gal, and that for the optimistic and pessimistic cases is $5.01/gal and $7.08/gal, 

respectively. The PV for the most probable, optimistic, and pessimistic cases is 69%, 47%, and 

108% more, respectively, than the PV estimated for the nth plant. The cost growth analysis shows 

that the TCI and Lang factor increased significantly from the base case nth plant. For the most 

probable cost growth scenario, the TCI and Lang factor are MM$886.4 and 8.11, respectively, 

which are an increase of 136% from the base case nth plant values.

Table 5. Pioneer Plant Analysis Results for the Dilute Acid Pretreatment Process Scenario

Cost Item Cost Growth (Pioneer Plant)

Most Probable Optimistic Pessimistic

PV ($/gal) 5.76 5.01 7.08

Fixed Capital 

Investment (MM$) 833 674 1,111

Total Capital 

Investment (MM$) 886 727 1,164

Lang Factor 8.11 6.65 10.65
25

Comparison with Previous Studies

The results of this study deviate considerably from a number of previous techno-economic 

analyses of cellulosic ethanol production. There are many contributing factors to this deviation,

and an explanation of the most significant of these factors is discussed here. Figure 14 presents a 

plot of estimated ethanol prices from seven previous studies as a function of feedstock price. The 

ethanol and feedstock prices were updated to 2007 dollars using the Consumer Price Index. The 

solid line on the plot represents the PV for the dilute acid pretreatment scenario using the model 

developed in this study as a function of feedstock price. 

Figure 14. Ethanol cost estimations from previous techno-economic studies.

(1) Short term technology—Simultaneous saccharification and fermentation (SSF). (2) Middle term 

technology—Simultaneous saccharification and cofermentation (SSCF). (3) Long term technology—

Consolidated bioprocessing (CBP). (4) Separate hydrolysis and fermentation (SHF). (5) SSF. (6) CBP. (7) 

SSCF.

After updating the feedstock and ethanol prices to 2007 dollars, much of the difference from 

previous studies can be explained by the clear correlation that exists between feedstock price and 

ethanol price. However, all of the studies except that of Nguyen and Saddler [22] remain lower 

than the line derived from this study. The study by Hamelinck et al. [23] represents a significant 

outlier from the apparent correlation between feedstock price and ethanol price. The three 

ethanol price estimates are for short-term (5 years from time of study), middle-term (10–15 

years), and long-term (20+ years) technology implementation. The short-term estimate is closest

to the time frame considered in this study. However, it also deviates from the trend of other 

studies. The assumptions for the short-term estimate—including feedstock input, rate of return, 

and reaction conversions—are quite similar to those used in this study, and the TCI (updated to 

2007 dollars) is nearly equal as well. The most significant difference from this study is the nonfeedstock operating cost, which is approximately $0.32/gal ethanol (EtOH) compared to 
26

$1.68/gal. This is partly due to lower costs for corn steep liquor, cellulase, and other raw 

materials. This factor accounts for most of the discrepancy between ethanol price estimates. 

The ethanol price from the study published by Sendich et al. [24] is also slightly lower than the 

apparent correlation of feedstock and ethanol price. The lowest estimate in that study assumes 

the use of consolidated bioprocessing, which is an advanced technology also modeled in the 

long-term estimate from Hamelinck et al. [23]. The higher ethanol price estimate of $1.03/gal is 

from a model using simultaneous saccharification and cofermentation (SSCF). SSCF is a more 

advanced technology than was considered in this study; this factor results in lower capital and 

operating costs by combining enzymatic saccharification and fermentation. A new AFEX 

pretreatment scheme was also employed, which may have contributed to lower capital and 

operating costs of pretreatment. 

The enzyme cost used in this study is much higher than that used in other studies. Because 

enzyme cost is such a significant fraction of the PV, it contributes significantly to the 

discrepancy between the current study and previous studies. For example, the enzyme prices 

used in the prior studies of Wingren et al. (2003), Wingren et al. (2004), and Aden et al. [25, 26,

4] are approximately 30%, 30%, and 17% of the price used in this study, respectively. 
27

Conclusions

The present study is based on published technological and economical data that in many ways

lacks specifics and details. Thus, a list of assumptions was developed. Sensitivity analysis on 

those assumptions was performed, and their impact on TCI and PV was wide. The studies 

identify the strengths and weaknesses of the technology. The major strength of the technology is 

its pre-commercial maturity. Several pilot-scale cellulosic ethanol plants are in development 

today. The U.S. Department of Energy has funded 10 pioneer plant projects producing biofuels 

using a biochemical pathway [27]. However, technological and economic data from these pilot 

and demo plants are not publically available for use in this study. 

The published data in 2007 shows that the technology can reach 4.5% of ethanol concentration in 

the fermented beer, which is approximately one-third of what commercial grain ethanol plants 

are achieving. The enzyme and feedstock costs are two major cost contributors. In our studies,

cellulase price is assumed to be nearly $0.70/gal of ethanol produced and its loading is 18.8 

FPU/g of cellulose. The high cost of enzymes represents a significant opportunity to reduce the 

PV through improved biotechnology. The current feedstock cost is assumed to be $83/MT. 

However, the cost is likely to be location sensitive. It is important to critically estimate the 

feedstock cost, which may limit the plant locations to certain areas. Published improvements in 

process technology since 2007 have not been considered in this work, and these improvements 

would likely decrease PV for a project being designed today. 
28

References 

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Technologies.” Bioresource Technology; Vol. 96(18), 2005; pp. 2019-2025.

2. Elander, R.; Eggeman, T. Alternative Pretreatment (CAFI)—Inital Data And Economic 

Analysis. Report. Golden, CO: National Renewable Energy Laboratory, 2004.

3. Wyman, C.E.; Dale, B.E.; Elander, R.T.; Holtzapple, M.; Ladish, M.R.; Lee, Y.Y. 

“Comparative Sugar Recovery Data from Laboratory Scale Application of Leading 

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2026-2032.

4. Aden, A.; Ruth, M; Ibsen, K.; Jechura, J.; Neeves, K.; Sheehan, J.; Wallace, B.; 

Montague, L.; Slayton, A.; Likes, 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, 2002.

5. Neeves, K.; Ibsen, K. Lignocellulosic Biomass-to-Ethanol Process Design and 

Economics Using Two-Stage Dilute Acid Hydrolysis. Report. Golden, CO: National 

Renewable Energy Laboratory, 2000.

6. Nguyen, Q.A.; Tucker, M.P.; Keller, F.A.; Eddy, F.P. “Two-Stage Dilute-Acid 

Pretreatment of Softwoods.” Applied Biochemistry and Biotechnology; Vol. 84-86, 2000;

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Schedule Slippate and Startup Costs. Report No. R-3215-PSSP/RC. Santa Monica, CA: 

RAND Corp., June 1986.

9. Aden, A. Biochemical Production of Ethanol from Corn Stover: 2007 State of 

Technology Model. Report No. NREL/TP-510-43205. Golden, CO: National Renewable 

Energy Laboratory, 2008.

10. Teymouri, F.; Laureano-Perez, L.; Alizadeh, H.; Dale, B.E. “Optimization of the 

Ammonia Fiber Explosion (AFEX) Treatment Parameters for Enzymatic Hydrolysis of 

Corn Stover.” Bioresource Technology; 96, 2005; pp. 2014-2018.

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Conversion Process Configurations for the Production of Cellulosic Ethanol.” 

Biotechnology Progress; Vol. 26(1), 2009; pp. 64-72. DOI: 10.1002/btpr.311.
29

12. Hsu, D. Process Economics of Pervaporation Membrane Separation of Ethanol and 

Water in the Biochemical Design Report. Technical Memo. Golden, CO: National 

Renewable Energy Laboratory, 2008.

13. O’Brien, D.J.; Roth, L.H.; McAloon, A.J. “Ethanol Production by Continuous 

Fermentation-Pervaporation: A Preliminary Economic Analysis.” Journal of Membrane 

Science; Vol. 166(1), 2000; pp. 105-111.

14. Wooley, R.; Ruth, M.; Sheehan, J.; Ibsen, K.; Majdeski, H.; Galvez, A. Lignocellulosic 

Biomass to Ethanol Process Design and Economics Utilizing Co-current Dilute Acid 

Prehydrolysis and Enzymatic Hydrolysis Current and Futuristic Scenarios. Report No. 

NREL/TP-580-26157. Golden, CO: National Renewable Energy Laboratory, 1999.

15. Merrow, W.E.; Phillips, K.E.; Myers, C.W. (1981). Understanding Cost Growth and 

Performance Shortfalls in Pioneer Process Plants. Report No. RAND/R-2569-DOE. 

Santa Monica, CA: The Rand Corporation, 1981.

16. Wallace, B. Cost Estimation for a Lignocellulosic Biomass to Ethanol Pioneer Plant. 

NREL Technical Memo. Golden, CO: National Renewable Energy Laboratory, 2003.

17. “Chemical Engineering Plant Cost Index.” Chemical Engineering. November 2008.

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Engineers. Fifth Edition. New York: McGraw Hill, 2004.

19. “Inorganic Chemical Index.” SRI International Economics Handbook, Economic 

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20. “Labor Index.” Washington, DC: U.S. Bureau of Labor Statistics, 2008.

21. Bryant, C. “Step Change in Cellulosic Ethanol – The Future is Moving Closer.”

Presented at the National Ethanol Conference, San Antonio, TX, 2009.

22. Nguyen, Q.A.; Saddler, J.N. “An Integrated Model for The Technical and Economic 

Evaluation of an Enzymatic Biomass Conversion Process.” Bioresource Technology;

Vol. 35(3), 1991; pp. 275-282. 

23. Hamelinck, C.N.; van Hooijdonk, G.; Faaij, A.P.C. “Ethanol from Lignocellulosic 

Biomass: Techno-Economic Performance in Short-, Middle- and Long-Term.” Biomass 

and Bioenergy; Vol. 28, 2005; pp. 384-410.

24. 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, 2009; pp. 8429-8435.
30

25. Wingren, A.; Galbe, M.; Zacchi, G. “Techno-Economic Evaluation of Producing Ethanol 

from Softwood: Comparison of SSF and SHF and Identification of Bottlenecks.”

Biotechnology Progress; Vol. 19, 2003; pp. 1109-1117.

26. Wingren, A.; Soderstrom, J.; Galbe, M.; Zacchi, G. “Process Considerations and 

Economic Evaluation of Two-Step Steam Pretreatment for Production of Fuel Ethanol 

from Softwood.” Biotechnology Progress; Vol. 20, 2004; pp. 1421-1429.

27. “Major DOE Biofuels Project Locations.” Washington, DC: U.S. Department of Energy, 

2008. Available at 

http://www.energy.gov/media/Major_DOE_Biofuels_Project_Locations.pdf; accessed 

January 27, 2010.
31

Appendices

Appendix A - Assumptions for Techno-economic Studies of Biochemical 

Conversion Processes

Plant Size, Location, and Construction

• Optimum plant size is regarded as economically feasible plant size, which would be a 

plant with capacity of 2000 MT/day (dry feedstock).

• The plant produces EtOH (as product) and electricity (as byproduct).

• The plant is considered to be located in the middle of corn farmland.

• 25% of the land will be tied up in infrastructure (roads and buildings).

• 75% of the farm land plants corn. 

• The plant will be designed based on the state of the technology (as of 2007), and it would 

be the nth plant of its kind. RAND/risk analysis will inflate price from mature nth plant to 

immature pioneer plant.

• The online time would be 350 days per year (equivalent capacity factor of 96%). 

• Construction time of 24 months is considered.

• Startup period would be 25% of the construction time (6 months).

• During this period, an average of 50% production will be achieved with expenditure of 

about 75% of variable expenses and 100% of fixed expenses. 

Units 

• Plant capacity, mass flow, and yields are based on metric tons (MT) per day.

• Feedstock cost and purchased chemicals costs are based on short tons (ST).

• Ethanol sale price is in $/gallon, and byproduct credit (electricity) is in $/kWh. 

• Process operating conditions are:

o Temperature in oC

o Pressure in bar

o Mass flow rates in MT/day. 

Feedstock and Enzymes

• Corn stover (composed of stalks, leaves, cobs, and husks) is considered as feedstock.

o The feedstock will be delivered to the feed handling area of the plant.

o Moisture content in the feedstock is 25% (wet basis).

o Feed composition is assumed to be the same as CAFI II analysis results. Table A1 displays feedstock composition.

o Feedstock is delivered in bales.

o The feedstock transportation and management protocol are not considered.
32 

o Feed cost is assumed to be $83/MT ($75/dry short ton) at the gate. 

• 72 hours of on-site storage (corn stover bales) is considered, and long term storage costs 

are not included in the analysis. 

• Stover is washed of dirt and metal is removed; no biomass is lost in washing. 

• Enzyme will be purchased or produced on-site through purchase/licensing agreements 

with enzyme suppliers.

o Enzyme loading is assumed as 31.3 mg protein/g cellulose in original feed. 

o Enzyme cost is $507/MT broth when purchased from off-site sources.

Table A-1. Corn Stover Feedstock Composition

Components Composition 

(%) Components Composition (%)

Extractives 8.26 Lignin 10.69

Cellulose 33.43 Ash 5.93

Xylan 22.16 Acetate 5.44

Galactan 1.36 Protein 2.24

Arabinan 4.08 Soluble 

Solids 5.83

Mannan 0.58 Moisture 25

Material and Energy Balance

Material Balance 

• Reactions and conversions of hemicellulose carbohydrates (arabinan, mannan, galactan) 

are assumed to have the same value as xylan in the pretreatment hydrolyzer (depends on 

pretreatment). 

• A total of 7% of fermenting sugars is assumed to be lost to contamination, which may be 

regarded as bad batches (the bad batches are dumped and regarded as loss). 

• The total amount of water from the pressure filter (bottom product of the first distillation 

column) that is directly recycled is set to be 25% (to minimize contaminant buildup in the 

stream). 

• Boiler blow down is considered to be 3% of steam production. 

• Carbon efficiency can be calculated based on carbohydrate carbon content, as follows:

• Cooling tower windage is 0.1% of the total flow to the tower.

o The tower blow down is 10% of the of the sum of the evaporative loss plus 

windage. 

• Well water will be used as process makeup water (lost to evaporation, blow down, 

windage, in solid waste).

ൌ ሻ%ሺ ݕ݂݂݁݅ܿ݅݁݊ܿ ݊݋ܾݎܽܥ

݈݋݄ܽ݊ݐܧ ݅݊ ݊݋ܾݎܽܥ

100 ൈ ݁ݐܽݎ݀ݕ݄݋ܾݎܽܥ ݏݏ݉ܽ݋݅ܤ ݅݊ ݊݋ܾݎܽܥ
33 

Energy Balance 

• Heat loss from the reactor will not be accounted for in energy balance calculations.

• Heat loss from the combustor is accounted for, totaling about 2.2%. 

• Electricity will be generated by burning lignin and waste (process waste and pressed 

solids from wastewater treatment), which will be used in the plant, and the net surplus 

will be sold to grid at a price of $0.054/kW (2007 dollar value). 

Equipment Design, Material of Construction, and Costing

Equipment Design

• The reactors will be modeled using experimentally determined conversions of specific 

reactions (kinetic expressions will be not used because of the level of their development).

• If the size of any equipment is known to change linearly with the inlet flow, that 

information can be used for equipment scaling (a characteristic of the size might be the 

heat duty for a heat exchanger if the log-mean temperature difference is known not to 

change). 

• For some equipment, nothing can be easily related to the size, in which case the unit will 

be resized with each process change (for example heat exchangers with varying 

temperature profiles; in this case, the heat exchanger area will be calculated each time the 

model will be run and the cost will be scaled using the ratio of the new and original 

areas).

Material of Construction

• The materials of construction for all equipment (except the pretreatment reactor, flash 

tank, and Pneumapress equipment) will be as follows (Delta-T/NREL experience):

o SS316 for flash tank (for the solid-liquid separation equipment) 

o Incoloy 825-clad steel for continuous pretreatment reactor and its parts in contact 

with acid, and SS316L for most other parts including presteamer.

Costing

• All pumps, tanks, screw conveyors, agitators, heat exchangers, and surge tanks will be 

estimated using ICARUS Process Evaluator and NREL database. 

• Large vessels (saccharification tank, fermentors, seed hold tanks), boiler feed water 

softening equipment, anaerobic and aerobic digesters, and filter press will be estimated 

from quotation. 

• Smaller vessels (seed fermentor) and coil coolers will be estimated by ICARUS and 

NREL database.

• If process changes are made and the equipment size changes, the equipment will be recosted following the exponential scaling expression:

*or characteristic linearly related to the size

ܰ݁ݓ ݐݏ݋ܥ ൌ ܱݎ݈݅݃݅݊ܽ Cost ൬ New size כ

Original size כ൰

ୣ୶୮
34

o The scaling exponents are obtained from NREL’s vendor quote database.

• The purchased equipment cost obtained in a particular year (before 2007) will be indexed 

to the year of 2007 using the Chemical Engineering Index.

• Installation factors (from Delta-T/NREL experience for aqueous-based process) will be 

applied to purchased equipment costs to determine the installation cost (not ICARUS).

Chemical Costing 

• Cost of acids and other chemicals, including boiler feed water softening chemical cost,

will be obtained from quotation.

o The cost of the chemicals will also be indexed following Industrial Inorganic 

Chemical Index (from SRI) to estimate the cost of the chemicals in the year of 

2007.

Operating Cost

• Working capital is assumed to be 15% of fixed capital investment.

o It is assumed that the product will be made and shipped and payment received in 

30 days.

• Annual maintenance materials will be 2% of the total installed equipment cost.

• Employee salaries will be indexed, if required, to the year 2007 following the data of the 

Bureau of Labor Statistics.

• Salaries of the yard employees will not include benefits and will be covered in the 

general overhead category.

o General overhead will be a factor of 60% applied to the total salaries 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.

Wastewater Treatment Plant

• The process will be designed for zero discharge to a municipal treatment plant in a 

steady-state mode, and the treated water will be suitable for recycling to the process.

• Any process upset (sudden increase of solids in the wastewater) will not be considered in 

the model.

• Rain and snow run-off, equipment washing, and other non-process waters are assumed to 

flow to the municipal wastewater treatment system; other intermittent loads (process 

spills) will not be considered in the design.

• No insoluble components (cellulose, xylan) will be included in the chemical-oxygendemand (COD) calculations because of uncertainty of their reactivity.

• Biological oxygen demand in the anaerobic digester is assumed to be 70% of the COD. 

o COD reduction in both digesters is considered as 99.4%.
35

Greenhouse Emissions and Control 

• All of the sulfur entering into the combustor is converted to SO2.

o 1% of the generated SO2 is converted to sulfuric acid.

o Flue gas temp will be kept above the dew point of sulfuric acid.

• Carbon monoxide is assumed to be generated at a rate of 0.31 kg/MWhr. 

o Unburned carbon (char) in the ash is low at 1%.

• NOx is generated at 0.31 kg/MWhr.

• Baghouse efficiency is taken as 98.8% (from Foster Wheeler Energy Limited experience) 

and will control the emission level below new source performance standard limit.

• The nitrogen level in the combined feed to the combustor is similar to coal when 

ammonia species are included, but more like untreated biomass when they are not.

o Impact of global warming potential of NOx is small relative to global warming 

potential of CO2 emission from the process.

Cost Analysis

• The total plant investment cost will be determined by applying overhead and contingency 

factors (NREL, ConocoPhillips Company experience and literature) to installed 

equipment costs.

o 20% (of total direct and indirect cost) contingency is assumed for nth plant, and 

30% for pioneer plant.

• Warehouse cost is 1.5% of total installed equipment cost, and site development cost is 

9% of the installed cost of process equipment areas (A100, A200, A300 and A500) 

(Delta-T/NREL/published data).

o The estimates are location sensitive.

• Total installed cost (TIC) includes total installed equipment cost, warehouse cost, and site 

development cost.

• Indirect cost involves:

o Engineering and supervision (32% of purchased equipment cost)

o Construction expenses (34% of purchased equipment cost)

o Legal and contractor’s fees (23% of purchased equipment cost). 

• Total direct and indirect cost (TD&IC) includes TIC and indirect costs.

• Working capital is assumed as 15% of FCI (FCI is the sum of TD&IC and contingency).

• Total capital investment is the sum of FCI and working capital.

• To determine the product value per gallon of ethanol (PV), a discounted cash flow 

analysis will be used (after knowing the major three costs areas: (i) total capital 

investment, (ii) variable operating costs, and (iii) fixed operating cost).
36

o A 10% discounted cash flow rate of return will be used over a 20-year plant life.

o The plant is considered 100% equity financed.

• For federal tax return purposes, depreciation will be determined as follows :

o IRS modified accelerated cost recovery system, which includes general 

depreciation system, will be followed that allows both the 200% and 150% 

declining balance (DB) methods of depreciation.

o This allows the shortest recovery period and the largest deductions.

o According to the IRS, the steam production plant should use a 20-year recovery 

period (depreciated over 20 years). 

o Any other property not specifically described in the publication should be 

depreciated using a 7-year recovery period. 

o Property listed with a recovery period less than 10 years will use the 

200% DB depreciation method and a 20-year recovery period property 

will use the 150% DB depreciation.

o State tax will not be considered for the calculation (because the location of the 

plant is not specified).

• Return on investment will be calculated on a per gallon basis. Income tax will be 

averaged over the plant life and that average will be calculated on a per gallon basis. 
37

Appendix B - Sensitivity Parameters and Values

Table B-1. Sensitivity Parameters for Pretreatment and Saccharification (AREA 200)

Parameters

Dilute Acid (2007 EVD) 2-Stage Dilute acid Hot Water AFEX

Base Case 

(50%)

Sensitivity 

(0-100%)

Base Case (50%) Sensitivity (0-100%) Base Case 

(50%)

Sensitivity 

(0-100%)

Base Case 

(50%)

Sensitivity 

1 (0-100%) st Stage 2nd Stage 1st Stage 2nd Stage

Pretreatment

Temperature (oC) 190 190-200 190 210 180 190 190-200 108 90-110

Pressure (atm) 11.6 12 18.5 12.6 12.6-15.4 18.7

Retention time (min) 2 1-10 2 15 15-5 5 5

Catalyst (Acid/NH3) 

conc (%) 1.9 0.71-2.4 1.64 1.80 0.5-2.66 2.6 0 N/A 1.1 0.5-1.3

Solid consistency (%) 29.6 18-29.6% 30 30 12.9 12.9-20 48 20-70

Conversion (%)

 Cellulose → Glucose 6.3 6.3-23 6.3 42.4 23 48 0.32 2-2.5 0 N/A

 Xylan → Xylose oligomer 2.7 2.7 0 0 N/A 0.5 N/A

 Xylan → Xylose 82.5 33-89.7 82.5 14.3 33-89.7 0-41.5 2.4 7.3-25 0 N/A

Saccharification

Conversion (%)

 Cellulose → Glucose 91.09 67-97 89.97 65 95.9 75

 Xylan → Xylose 57.13 52.4 56.61 63 77.7 55-90
38

Table B-2. Sensitivity Parameters for Overall Process

Parameters

Dilute Acid (2007 EVD) 2-Stage Dilute acid Hot Water AFEX

Base Case 

(50%)

Sensitivity 

(0-100%)

Base Case 

(50%)

Sensitivity 

(0-100%)

Base Case 

(50%)

Sensitivity 

(0-100%)

Base Case 

(50%)

Sensitivity 

(0-100%)

Feedstock Cost ($/MT)

a 83 55-110 55 83-110 83 55-110 83 55-110

Pretreatment Reactor Installed 

Cost (MM$)b 23 14-32.7 0.32 0.22-0.62 25.7 20-40

Project Contingency (%) 20 10-30 20 20-30 20 10-30 20 10-30

Installation Factor or Cost (MM$)

 Factor 2.58 2.50-3.05 2.59 2.50-3.05 2.05 2-4

 Total Installed Equipment Cost 164.1 159-194 156 151-184 211.5 200-275

Enzyme Cost ($/MT broth) 507 256-1460 507 256-1460 507 256-1460

Enzyme loading (mg protein/g

cellulose) 31.3 20-33.3 31.3 20-33.3 31.3 20-33.3 31.3 20-33.3

Value of Excess Electricity/ 

Byproduct Credit ($/kWh) 0.054 0.03-0.06 0.054 0.03-0.05 0.04 0.03-0.05 0.04 0.03-0.05

a Feedstock contains 25% moisture. b 2007 indexed installed equipment cost. 
39

Appendix C - Sensitivity Results

Table C-1. Impact of Pretreatment Parameters on PV

Sensitivity Production 

(MMGal/Yr)

PV

Parameter Values ($/Gal) Change 

(%)a

Dilute Acid Pretreatment — 2007 EVD Scenario 

Base Case 53.4 3.40

Pretreatment

Highb

Scenarios

Reactor 

temperature 

(

o

C)

200.0 53.4 3.40 0

Residence 

time (min) 10 53.4 3.93 15.6

Acid 

concentration 

(%)

2.4 53.4 3.40 0

Cellulose to 

Glucose 

(% conv)

23 53.4 3.42 0.6

Xyl to Xylose 

(% conv) 89.7 53.9 3.37 -0.9

Lowb

Scenarios

Residence 

time (min) 1 53.4 3.34 -1.8

Solid 

consistency 

(%)

18.0 53.4 3.47 2.1

Acid 

concentration 

(%)

0.71 53.4 3.40 0

Xyl to Xylose 

(% conv) 33 50.1 3.61 6.2

Saccharification

High 

Scenarios

Cellulose to 

Glucose 

(% conv)

97 54.9 3.31 -2.6

Low 

Scenarios

Cellulose to 

Glucose 

(% conv)

67 44.6 4.07 19.7

Xylan to 

Xylose (%

conv)

52.4 53.3 3.41 0.3

Two-Stage Dilute Acid Pretreatment — 2007 EVD Scenario

Base Case 42.5 4.38

1st Stage Acid 

Pretreatment 

(Pretreatment)

High 

Scenarios

Cellulose to 

Glucose 

(% conv)

23 47.3 3.93 -10.3

Xylan to 

Xylose 

(% conv)

89.7 44.4 4.19 -4.3
40

Sensitivity Production 

(MMGal/Yr)

PV

Parameter Values ($/Gal) Change 

(%)a

Other 

hemicellulose 

to monomers 

(% conv)

90 42.5 4.38 0

Acid 

Concentration 

(%)

2.66 42.5 4.38 0

Low 

Scenarios

Reactor 

temperature 

(

o

C)

180 42.5 4.37 -0.2

Acid 

concentration 

(%)

0.5 42.5 4.38 0

Xylan to 

Xylose 

(% conv)

33 29.5 6.29 43.6

2nd Stage Acid 

Pretreatment

(Saccharification)

High

Scenarios

Cellulose to 

glucose 

(% conv)

48 45.7 4.08 -6.8

Xylan to Xylose

(% conv) 41.5 43.9 4.24 -3.2

Acid 

concentration 2.6 42.4 4.41 0.7

Low 

Scenarios

Xylan to 

Xylose 

(% conv)

0 41.8 4.45 1.6

Hot Water Pretreatment — 2007 EVD Scenario

Base Case 39.0 4.29

Pretreatment

High 

Scenarios

Reactor 

temperature 

(

o

C)

200 39.0 4.29 0

Residence 

time (min) 20 39.0 4.29 0

Solid 

consistency 

(%)

20.0 39.0 3.84 -10.2

Cellulose to 

Glucose 

(% conv)

2 39.1 4.29 0

Xylan to 

Oligomer (% 

conv)

60 38.5 4.36 1.6

Xylan to 

Xylose (% 

conv)

7.3 39.5 4.24 -1.1

Low 

Scenarios

Residence 

time (min) 5 39.0 4.29 0

Xylan to 

oligomer 

(% conv)

25 42.9 3.92 -8.4
41

Sensitivity Production 

(MMGal/Yr)

PV

Parameter Values ($/Gal) Change 

(%)a

Xylan to 

Xylose 

(% conv)

63 39.6 4.23 -1.4

Saccharification

High 

Scenarios

Cell Cellulose 

to Glucose 

(% conv)

65 29.7 5.61 30.7

Low 

Scenarios

Cellulose to 

Glucose 

(% conv)

65 29.7 5.61 30.7

AFEX — 2007 EVD Scenario

Base Case 46.2 3.69

Pretreatment

High 

Scenarios

Reactor 

temperature 

(

o

C)

110 46.1 3.69 0

Solids 

consistency 

(%)

70 46.2 4.16 12.7

Xylan to 

Oligomer 

(%conv)

20 46.2 3.69 0

Ammonia 

loading (ratio

to dry stover)

1:0.8 46.2 3.69 0

Low 

Scenarios

Reactor 

temperature 

(

o

C)

80 46.0 3.68 -0.3

Solid 

consistency 

(%)

20 46.2 3.31 -10.3

Xylan to 

Xylose 

(%conv)

90 46.9 3.65 -1.1

Ammonia 

loading (ratio

to dry stover)

1:1.3 46.2 3.70 0.3

Saccharification

High 

Scenarios

Xylan to 

Xylose 

(%conv)

55 45.8 3.80 3.0

Low 

Scenarios

Cellulose to 

Glucose 

(%conv)

75 43.2 3.87 4.9

a Percentage change of PV from 2007 EVD scenarios. b High and low refer to the value of each parameter with 

respect to the base case. Not all parameters were tested at both high and low scenarios. 
42

Table C-2. Impact of Overall Process and Economic Parameters on PV

Sensitivity Parameters

Base Case Low High

Value PV

($/gal) Value PV($/gal) Value PV($/gal)

Feedstock Cost ($/MT) 83 3.40 55 3.04 110 3.77

Pretreatment Reactor (MM$) 23 3.40 14 3.34 32.7 3.47

Installed Equipment Cost (MM$) 164.1 3.40 159 3.38 194.0 3.55

(Corresponding Installation factor) (2.58) (2.5) (3.05)

Enzyme Cost ($/MT of broth) 507 3.40 256 3.05 1460 4.73

Electricity Price ($/kWh) 0.054 3.40 0.06 3.38 0.03 3.50

Enzyme Loading (mg/g protein) 31.3 3.40 20 3.15 33.3 3.45

Contingency (%) 20 3.40 10 3.31 30 3.50
43

Appendix D - Cost Analysis Result Summary 

Table D-1. Cost Analysis Result Summary for Dilute Acid Pretreatment Processes

Ethanol Production Process Engineering Analysis

UCR Dilute Acid - Corn Stover, Current Case

Dilute Acid Prehydrolysis with Saccharification and Cofermentation

All Values in 2007$

Product Value $3.40

Ethanol Production (MM Gal. / Year) 53.4 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 76.3

Feedstock Cost $/Dry MetricTon $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

Feed Handling $10,900,000 Feedstock 108.4

Pretreatment $36,200,000 Biomass to Boiler 0.0

Saccharification & Fermentation $21,800,000 CSL 16.0

Distillation and Solids Recovery $26,100,000 Cellulase 69.5

Wastewater Treatment $3,500,000 Other Raw Materials 17.8

Storage $3,200,000 Waste Disposal 12.7

Boiler/Turbogenerator $56,100,000 Electricity -21.9

Utilities $6,300,000 Fixed Costs 18.5

Total Installed Equipment Cost $164,100,000 Capital Depreciation 30.5

Average Income Tax 26.7

Added Costsa $211,800,000 Average Return on Investment 62.2

 (% of TCI) 56% Operating Costs ($/yr)

Working Capital 49,030,000 Feedstock $57,900,000

Total Capital Investment $375,900,000 Biomass to Boiler $0

CSL $8,500,000

Installed Equipment Cost/Annual Gallon $3.07 Cellulase $37,100,000

Total Project Investment/Annual Gallon $7.04 Other Raw Matl. Costs $9,500,000

Waste Disposal $6,800,000

Loan Rate N/A Electricity -$11,700,000

Term (years) N/A Fixed Costs $9,900,000

Capital Charge Factor 0.170 Capital Depreciation $16,300,000

Average Income Tax $14,300,000

Denatured Fuel Prod. (MMgal / yr) 55.9 Average Return on Investment $33,300,000

Denatured Fuel Min. Sales Price $3.29 Excess Electricity (KWH/gal) 4.06

Denaturant Cost ($/gal denaturant) $0.739 Plant Electricity Use (KWH/gal) 2.23

Plant Steam Use (kg steam/gal) 17.5

Maximum Yields (100% of Theoretical) Boiler Feed–LHV (Btu/lb) 2,209

 Ethanol Production (MM Gal/yr) 82.5 Boiler Feed–Water Fraction 0.542

 Theoretical Yield (Gal/ton) 106.9 Specific Operating Conditions

Current Yield (Actual/Theoretical) 65% Enzyme Loading (mg/g cellulose) 33.5

Saccharification Time (days) 5.0

Conversion Cellulose-->Glucose 0.9109

Fermentation Time (days) 2.0 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
44

Table D-2. Cost Analysis Result Summary for Dilute Acid Pretreatment (High Solids) Processes

a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.

Ethanol Production Process Engineering Analysis

UCR Dilute Acid - Pretreatment yields based on NREL FY08 SOT

Dilute Acid Prehydrolysis with Saccharification and Cofermentation

All Values in 2007$

Product Value $3.60

Ethanol Production (MM Gal. / Year) 50.8 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 72.5

Feedstock Cost $/Dry Metric Ton $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

Feed Handling $10,900,000 Feedstock 114.0

Pretreatment $38,000,000 Biomass to Boiler 0.0

Saccharification & Fermentation $21,800,000 CSL 16.5

Distillation and Solids Recovery $25,700,000 Cellulase 73.1

Wastewater Treatment $5,800,000 Other Raw Materials 18.8

Storage $3,100,000 Waste Disposal 13.0

Boiler/Turbogenerator $57,600,000 Electricity -24.8

Utilities $6,600,000 Fixed Costs 19.8

Total Installed Equipment Cost $169,400,000 Capital Depreciation 33.3

Average Income Tax 29.0

Added Costsa $219,100,000 Average Return on Investment 67.5

 (% of TCI) 56% Operating Costs ($/yr)

Working Capital 50,670,000 Feedstock $57,900,000

Total Capital Investment $388,500,000 Biomass to Boiler $0

CSL $8,400,000

Installed Equipment Cost/Annual Gallon $3.34 Cellulase $37,100,000

Total Project Investment/Annual Gallon $7.65 Other Raw Matl. Costs $9,600,000

Waste Disposal $6,600,000

Loan Rate N/A Electricity -$12,600,000

Term (years) N/A Fixed Costs $10,100,000

Capital Charge Factor 0.169 Capital Depreciation $16,900,000

Average Income Tax $14,700,000

Denatured Fuel Prod. (MMgal / yr) 53.1 Average Return on Investment $34,200,000

Denatured Fuel Min. Sales Price $3.47

Denaturant Cost ($/gal denaturant) $0.739 Excess Electricity (KWH/gal) 4.58

Plant Electricity Use (KWH/gal) 2.39

Maximum Yields (100% of Theoretical)

 Ethanol Production (MM Gal/yr) 82.5 Plant Steam Use (kg steam/gal) 18.2

 Theoretical Yield (Gal/ton) 106.9 Boiler Feed -- LHV (Btu/lb) 2,286

Current Yield (Actual/Theoretical) 62% Boiler Feed -- Water Fraction 0.539

Specific Operating Conditions

Enzyme Loading (mg/g cellulose) 35.1

Saccharification Time (days) 5.0

Conversion Cellulose --> Glucose 0.9109

Fermentation Time (days) 2.0
45

Table D-3. Cost Analysis Result Summary for Hot Water Pretreatment Process

Ethanol Production Process Engineering Analysis

Hot Water - Corn Stover, Current Case

Hot Water Prehydrolysis with Saccharification and Cofermentation

All Values in 2007$

Product Value $4.44

Ethanol Production (MM Gal. / Year) 39.0 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 55.7

Feedstock Cost $/Dry MetricTon $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

 Feed Handling $10,900,000 Feedstock 148.4

 Pretreatment $6,700,000 Biomass to Boiler 0.0

 Saccharification & Fermentation $30,200,000 CSL 28.5

 Distillation and Solids Recovery $30,900,000 Cellulase 95.1

 Wastewater Treatment $1,900,000 Other Raw Materials 5.1

 Storage $3,300,000 Waste Disposal 3.5

 Boiler/Turbogenerator $65,800,000 Electricity -29.0

 Utilities $6,700,000 Fixed Costs 24.5

Total Installed Equipment Cost $156,300,000 Capital Depreciation 40.2

Average Income Tax 36.3

Added Costsa $204,800,000 Average Return on Investment 90.9

 (% of TCI) 57% Operating Costs ($/yr)

Working Capital 47,100,000 Feedstock $57,900,000

Total Capital Investment $361,100,000 Biomass to Boiler $0

CSL $11,100,000

Installed Equipment Cost/Annual Gallon $4.01 Cellulase $37,100,000

Total Project Investment/Annual Gallon $9.26 Other Raw Matl. Costs $2,000,000

Waste Disposal $1,400,000

Loan Rate N/A Electricity -$11,300,000

Term (years) N/A Fixed Costs $9,600,000

Capital Charge Factor 0.181 Capital Depreciation $15,700,000

Average Income Tax $14,200,000

Denatured Fuel Prod. (MMgal / yr) 40.8 Average Return on Investment $35,500,000

Denatured Fuel Min. Sales Price $4.27

Denaturant Cost ($/gal denaturant) $0.739 Excess Electricity (KWH/gal) 5.37

Plant Electricity Use (KWH/gal) 3.30

Maximum Yields (100% of Theoretical)

 Ethanol Production (MM Gal/yr) 82.5 Plant Steam Use (kg steam/gal) 40.7

 Theoretical Yield (Gal/ton) 106.9 Boiler Feed -- LHV (Btu/lb) 2,354

Current Yield (Actual/Theoretical) 47% Boiler Feed -- Water Fraction 0.529

Specific Operating Conditions

Enzyme Loading (mg/g cellulose) 33.1

Saccharification Time (days) 5.0

Conversion Cellulose --> Glucose 0.8997

Fermentation Time (days) 2.0 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
46

Table D-4. Cost Analysis Result Summary for Two-Stage Dilute Acid Pretreatment Processes

Ethanol Production Process Engineering Analysis

Corn Stover Design Report Case: 2020 Market Target Case (2002 Design Report Target Case)

Dilute Acid Prehydrolysis and Hydrolysis

All Values in 2007$

Product Value $4.38

Ethanol Production (MM Gal. / Year) 32.8 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 46.9

Feedstock Cost $/Dry Metric Ton $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

 Feed Handling $10,900,000 Feedstock 176.4

 Pretreatment $44,900,000 Biomass to Boiler 0.0

 Saccharification & Fermentation $9,700,000 CSL 26.2

 Distillation and Solids Recovery $26,700,000 Cellulase 0.0

 Wastewater Treatment $4,600,000 Other Raw Materials 32.3

 Storage $2,800,000 Waste Disposal 22.0

 Boiler/Turbogenerator $66,200,000 Electricity -51.3

 Utilities $6,900,000 Fixed Costs 31.0

Total Installed Equipment Cost $172,700,000 Capital Depreciation 51.8

Average Income Tax 45.8

Added Costsa $218,300,000 Average Return on Investment 103.5

 (% of TCI) 56% Operating Costs ($/yr)

Working Capital (WC) 51,000,000 Feedstock $57,900,000

Total Capital Investment $391,000,000 Biomass to Boiler $0

CSL $8,600,000

Installed Equipment Cost/Annual Gallon $5.26 Cellulase $0

Total Project Investment/Annual Gallon $11.92 Other Raw Matl. Costs $10,600,000

Waste Disposal $7,200,000

Loan Rate N/A Electricity -$16,800,000

Term (years) N/A Fixed Costs $10,200,000

Capital Charge Factor 0.169 Capital Depreciation $17,000,000

Average Income Tax $15,000,000

Denatured Fuel Prod. (MMgal / yr) 34.3 Average Return on Investment $34,000,000

Denatured Fuel Min. Sales Price $4.21

Denaturant Cost ($/gal denaturant) $0.739 Excess Electricity (KWH/gal) 9.50

Plant Electricity Use (KWH/gal) 3.46

Plant Steam Use (kg steam/gal) 35.4

Maximum Yields (100% of Theoretical) Boiler Feed -- LHV (Btu/lb) 2,379

 Ethanol Production (MM Gal/yr) 82.5 Boiler Feed -- Water Fraction 0.521

 Theoretical Yield (Gal/ton) 106.9 Specific Operating Conditions

Current Yield (Actual/Theoretical) 40% Feed rate (dry tonnes/day) 2,000

Feed rate (dry tons/day) 2,205

Lignin Residue (dry tonnes/day) 979

Lignin Residue (dry tons/day) 1,079

Pretreatment Total Solids (wt%) 30

Saccharification Total Solids (wt%) 34 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
47

Table D-5. Cost Analysis Result Summary for AFEX Pretreatment Processes

Ethanol Production Process Engineering Analysis

Corn Stover Design Report Case: 2020 Market Target Case (2002 Design Report Target Case)

AFEX with Saccharification and Cofermentation

All Values in 2007$

Product Value $3.69

Ethanol Production (MM Gal. / Year) 46.2 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 66.0

Feedstock Cost $/Dry Metric Ton $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

 Feed Handling $10,900,000 Feedstock 125.3

 Pretreatment $30,800,000 Biomass to Boiler 0.0

 CSL 19.4

 Saccharification & Fermentation $23,500,000 Cellulase 80.4

 Distillation and Solids Recovery $27,500,000 Other Raw Materials 14.1

 Wastewater Treatment $1,600,000 Waste Disposal 2.9

 Storage $2,800,000 Electricity -36.7

 Boiler/Turbogenerator $62,000,000 Fixed Costs 21.6

 Utilities $8,400,000 Capital Depreciation 36.4

Total Installed Equipment Cost $167,400,000 Average Income Tax 32.0

Average Return on Investment 73.4

Added Costsa $218,600,000 Operating Costs ($/yr)

 (% of TCI) 57% Feedstock $57,900,000

Working Capital 50,350,000 Biomass to Boiler $0

Total Capital Investment $386,000,000 CSL $8,900,000

Cellulase $37,100,000

Installed Equipment Cost/Annual Gallon $3.63 Other Raw Matl. Costs $6,500,000

Total Project Investment/Annual Gallon $8.36 Waste Disposal $1,400,000

Electricity -$16,900,000

Loan Rate N/A Fixed Costs $10,000,000

Term (years) N/A Capital Depreciation $16,800,000

Capital Charge Factor 0.170 Average Income Tax $14,800,000

Average Return on Investment $33,900,000

Denatured Fuel Prod. (MMgal / yr) 48.3 Excess Electricity (KWH/gal) 5.23

Denatured Fuel Min. Sales Price $3.56 Plant Electricity Use (KWH/gal) 2.57

Denaturant Cost ($/gal denaturant) $0.739 Plant Steam Use (kg steam/gal) 35.9

Boiler Feed -- LHV (Btu/lb) 2,295

Maximum Yields (100% of Theoretical) Boiler Feed -- Water Fraction 0.538

 Ethanol Production (MM Gal/yr) 82.5 Specific Operating Conditions

 Theoretical Yield (Gal/ton) 106.9 Feed rate (dry tonnes/day) 2,000

Current Yield (Actual/Theoretical) 56% Feed rate (dry tons/day) 2,205

Lignin Residue (dry tonnes/day) 730

Lignin Residue (dry tons/day) 804

Pretreatment Total Solids (wt%) 48

Saccharification Total Solids (wt%) 20 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
48

Table D-6. Cost Analysis Result Summary for On-site Enzyme Production Processes

Ethanol Production Process Engineering Analysis

UCR Dilute Acid - Corn Stover, Current Case

Dilute Acid Prehydrolysis with Saccharification and Cofermentation (On-site Enzyme Production)

All Values in 2007$

Product Value $3.54

Ethanol Production (MM Gal. / Year) 47.4 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 67.7

Feedstock Cost $/Dry Metric Ton $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

 Feed Handling $10,900,000 Feedstock 122.1

 Pretreatment $36,200,000 Biomass to Boiler 0.0

 On-site Enzyme Production $23,700,000 CSL 19.8

 Saccharification & Fermentation $21,800,000 Cellulase 0.0

 Distillation and Solids Recovery $25,400,000 Other Raw Materials 20.4

 Wastewater Treatment $3,500,000 Waste Disposal 14.3

 Storage $2,500,000 Electricity 1.6

 Boiler/Turbogenerator $57,200,000 Fixed Costs 22.7

 Utilities $6,800,000 Capital Depreciation 39.9

Total Installed Equipment Cost $187,800,000 Average Income Tax 34.0

Average Return on Investment 79.2

Added Costsa $246,400,000 Operating Costs ($/yr)

 (% of TCI) 57% Feedstock $57,900,000

Working Capital 56,630,000 Biomass to Boiler $0

Total Capital Investment $434,200,000 CSL $9,400,000

Cellulase $0

Installed Equipment Cost/Annual Gallon $3.96 Other Raw Matl. Costs $9,700,000

Total Project Investment/Annual Gallon $9.16 Waste Disposal $6,800,000

Electricity $800,000

Loan Rate N/A Fixed Costs $10,700,000

Term (years) N/A Capital Depreciation $18,900,000

Capital Charge Factor 0.167 Average Income Tax $16,100,000

Average Return on Investment $37,500,000

Denatured Fuel Prod. (MMgal / yr) 49.6

Denatured Fuel Min. Sales Price $3.42 Excess Electricity (KWH/gal) -0.30

Denaturant Cost ($/gal denaturant) $0.739 Plant Electricity Use (KWH/gal) 7.63

Plant Steam Use (kg steam/gal) 19.7

Boiler Feed -- LHV (Btu/lb) 2,239

Maximum Yields (100% of Theoretical) Boiler Feed -- Water Fraction 0.538

 Ethanol Production (MM Gal/yr) Specific Operating Conditions

 Theoretical Yield (Gal/ton) 82.5 Enzyme Loading (mg/g cellulose) 0.0

Current Yield (Actual/Theoretical) 106.9 Saccharification Time (days) 5.0

57% Conversion Cellulose --> Glucose 0.9109

Fermentation Time (days) 2.0 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
49

Table D-7. Cost Analysis Result Summary for Pervaporation Purification Processes

Ethanol Production Process Engineering Analysis

Corn Stover Design Report Case: 2020 Market Target Case (2002 Design Report Target Case)

Dilute Acid Prehydrolysis with Saccharification and Cofermentation (Pervaporation Separation)

All Values in 2007$

Product Value $3.75

Ethanol Production (MM Gal. / Year) 53.9 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 77.0

Feedstock Cost $/Dry MetricTon $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

 Feed Handling $10,900,000 Feedstock 107.5

 Pretreatment $36,200,000 Biomass to Boiler 0.0

 CSL 15.8

 Saccharification & Fermentation $21,900,000 Cellulase 68.9

 Distillation and Solids Recovery $70,800,000 Other Raw Materials 18.4

 Wastewater Treatment $3,300,000 Waste Disposal 12.6

 Storage $3,300,000 Electricity -25.3

 Boiler/Turbogenerator $55,900,000 Fixed Costs 21.4

 Utilities $7,000,000 Capital Depreciation 40.5

Total Installed Equipment Cost $209,200,000 Average Income Tax 34.1

Average Return on Investment 81.0

Added Costsa $226,500,000 Operating Costs ($/yr)

 (% of TCI) 52% Feedstock $57,900,000

Working Capital $65,350,000 Biomass to Boiler $0

Total Capital Investment $501,000,000 CSL $8,500,000

Cellulase $37,100,000

Installed Equipment Cost/Annual Gallon $3.88 Other Raw Matl. Costs $9,900,000

Total Project Investment/Annual Gallon $8.09 Waste Disposal $6,800,000

Electricity -$13,600,000

Loan Rate N/A Fixed Costs $11,500,000

Term (years) N/A Capital Depreciation $21,800,000

Capital Charge Factor 0.192 Average Income Tax $18,400,000

Average Return on Investment $43,600,000

Denatured Fuel Prod. (MMgal / yr) 56.4 Excess Electricity (KWH/gal) 4.68

Denatured Fuel Min. Sales Price $3.61 Plant Electricity Use (KWH/gal) 2.25

Denaturant Cost ($/gal denaturant) $0.739 Plant Steam Use (kg steam/gal) 11.9

Maximum Yields (100% of Theoretical) Boiler Feed -- LHV (Btu/lb) 2,193

 Ethanol Production (MM Gal/yr) 82.5 Boiler Feed -- Water Fraction 0.543

 Theoretical Yield (Gal/ton) 106.9 Specific Operating Conditions

Current Yield (Actual/Theoretical) 65% Feed rate (dry tonnes/day) 2,000

Feed rate (dry tons/day) 2,205

Lignin Residue (dry tonnes/day) 626

Lignin Residue (dry tons/day) 690

Pretreatment Total Solids (wt%) 30

Saccharification Total Solids (wt%) 20 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
50

Table D-8. Cost Analysis Result Summary for Separate C5 and C6 Fermentation Processes

Ethanol Production Process Engineering Analysis

Corn Stover Design Case: 2005 Post Enzyme-Subcontract Case - 90% Cellulose in 7 days @ 1.9X loading

Dilute Acid Prehydrolysis with Saccharification and Cofermentation (Separate C5 & C6 Sugar fermentation)

All Values in 2007$

Product Value $3.67

Ethanol Production (MM Gal. / Year) 55.5 Ethanol at 68°F

Ethanol Yield (Gal / Dry Metric Ton Feedstock) 79.3

Feedstock Cost $/Dry MetricTon $83

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal ethanol)

 Feed Handling $11,400,000 Feedstock 104.3

 Pretreatment $34,200,000 Biomass to Boiler 0.0

 CSL 33.1

 Saccharification & Fermentation $29,600,000 Cellulase 66.9

 Distillation and Solids Recovery $27,300,000 Other Raw Materials 35.9

 Wastewater Treatment $3,100,000 Waste Disposal 2.6

 Storage $5,100,000 Electricity -11.6

 Boiler/Turbogenerator $50,900,000 Fixed Costs 18.0

 Utilities $6,200,000 Capital Depreciation 30.3

Total Installed Equipment Cost $167,800,000 Average Income Tax 26.0

Average Return on Investment 61.6

Added Costsa $218,000,000 Operating Costs ($/yr)

 (% of TCI) 57% Feedstock $57,900,000

Working Capital $50,320,000 Biomass to Boiler $0

Total Capital Investment $385,800,000 CSL $18,400,000

Cellulase $37,100,000

Installed Equipment Cost/Annual Gallon $3.02 Other Raw Matl. Costs $19,900,000

Total Project Investment/Annual Gallon $6.95 Waste Disposal $1,400,000

Electricity -$6,500,000

Loan Rate N/A Fixed Costs $10,000,000

Term (years) N/A Capital Depreciation $16,800,000

Capital Charge Factor 0.170 Average Income Tax $14,400,000

Average Return on Investment $34,200,000

Denatured Fuel Prod. (MMgal / yr) 58.1

Denatured Fuel Min. Sales Price $3.54 Excess Electricity (KWH/gal) 2.16

Denaturant Cost ($/gal denaturant) $0.739 Plant Electricity Use (KWH/gal) 2.33

Plant Steam Use (kg steam/gal) 19.9

Boiler Feed -- LHV (Btu/lb) 1,489

Maximum Yields (100% of Theoretical) Boiler Feed -- Water Fraction 0.547

 Ethanol Production (MM Gal/yr) 82.5 Specific Operating Conditions

 Theoretical Yield (Gal/ton) 106.9 Enzyme Loading (mg/g cellulose) 33.4

Current Yield (Actual/Theoretical) 67% Saccharification Time (days) 5.0

Conversion Cellulose --> Glucose 0.9109

Fermentation Time (days) 2.0 a Added costs include indirect costs (engineering and supervision, construction expenses, legal and contractor fees)

and contingency.
51

Appendix E - Equipment List and Costs, Installation Factors, and Installed Equipment Costs for Dilute Acid 

Pretreatment Processes

Table E-1. Equipment Lists and Costs for Dilute Acid Pretreatment Processes

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

C-101 2 Bale Transport 

Conveyor 98,040 111,111 1.13 $400,000 2000 $800,000 0.6 $862,388 1.62 $1,397,068 $1,862,521

C-102 2 Bale Unwrapping 

Conveyor 98,040 111,111 1.13 $150,000 2000 $300,000 0.6 $323,395 1.19 $384,840 $513,056

C-103 1 Belt Press Discharge 

Conveyor 98,040 111,111 1.13 $50,000 2000 $50,000 0.6 $53,899 1.89 $101,870 $135,809

C-104 4 Shredder Feed 

Conveyor 98,040 111,111 1.13 $60,000 2000 $240,000 0.6 $258,716 1.38 $357,028 $475,978

M-101 2 Truck Scales 98,040 111,111 1.13 $34,000 2000 $68,000 0.6 $73,303 2.47 $181,058 $241,380

M-102 4 1 Truck Unloading 

Forklift 98,040 111,111 1.13 $18,000 2000 $90,000 1 $101,999 1 $101,999 $135,982

M-103 4 Bale Moving Forklift 98,040 111,111 1.13 $18,000 2000 $72,000 1 $81,599 1 $81,599 $108,785

M-104 2 Corn Stover Wash 

Table 98,040 111,111 1.13 $104,000 2000 $208,000 0.6 $224,221 2.39 $535,888 $714,426

M-105 4 Shredder 98,040 111,111 1.13 $302,000 2000 $1,208,000 0.6 $1,302,205 1.38 $1,797,043 $2,395,754

M-106 1

Concrete 

Feedstock-Storage 

Slab

98,040 111,111 1.13 $450,655 2000 $450,655 1 $510,738 2.2 $1,123,624 $1,497,976

M-107 1 Polymer Feed 

System 98,040 111,111 1.13 $30,000 2000 $30,000 0.6 $32,340 2.28 $73,734 $98,300

P-101 2 1 Wash Table Pump 98,040 111,111 1.13 $20,000 2000 $60,000 0.79 $66,236 3.87 $256,332 $341,732

P-102 2 1 Wash Water Pump 98,040 111,111 1.13 $15,000 2000 $45,000 0.79 $49,677 5.19 $257,822 $343,719

P-103 1 1 Clarifier 

Underflow Pump 98,040 111,111 1.13 $6,000 2000 $12,000 0.79 $13,247 13.41 $177,644 $236,828

P-104 1 1 Clarified Water 

Pump 98,040 111,111 1.13 $15,000 2000 $30,000 0.79 $33,118 7.07 $234,143 $312,151

P-105 1 1 Belt Press Sump 

Pump 98,040 111,111 1.13 $19,000 2000 $38,000 0.79 $41,949 2.92 $122,492 $163,301

S-101 1 Clarifier Thickener 98,040 111,111 1.13 $135,000 2000 $135,000 0.6 $145,528 1.51 $219,747 $292,959
52

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

S-102 1 Belt Press 98,040 111,111 1.13 $100,000 2000 $100,000 0.6 $107,798 1.25 $134,748 $179,641

S-103 1 Magnetic 

Separator 159,948 111,111 0.69 $13,863 1998 $13,863 0.6 $11,141 1.3 $14,483 $19,537

T-101 1 Wash Water Tank 98,040 111,111 1.13 $50,000 2000 $50,000 0.51 $53,296 2.8 $149,227 $198,945

T-102 1 Clarifier Thickener 

Tank 98,040 111,111 1.13 $135,000 2000 $135,000 0.51 $143,898 3.04 $437,450 $583,192

A100 Subtotal $4,135,518 $4,490,691 1.81 $8,139,839 $10,851,970

A-201 1 In-line Sulfuric 

Acid Mixer 55,308 102,884 1.86 $1,900 1997 $1,900 0.48 $2,559 1 $2,559 $3,479

A-205 1 Hydrolyzate Mix 

Tank Agitator 358,810 236,880 0.66 $36,000 1997 $36,000 0.51 $29,129 1.2 $34,955 $47,517

A-209 1 Overliming Tank 

Agitator 167,050 301,831 1.81 $19,800 1997 $19,800 0.51 $26,773 1.3 $34,805 $47,313

A-224 1 Reacidification 

Tank Agitator 167,280 306,440 1.83 $65,200 1997 $65,200 0.51 $88,782 1.2 $106,539 $144,827

A-232 1 Reslurrying Tank 

Agitator 358,810 416,118 1.16 $36,000 1997 $36,000 0.51 $38,826 1.2 $46,591 $63,335

C-201 1 Hydrolyzate Screw 

Conveyor 225,140 236,880 1.05 $59,400 1997 $59,400 0.78 $61,802 1.3 $80,343 $109,217

C-202 1

Hydrolyzate 

Washed Solids 

Belt Conveyor

91,633 77,862 0.85 $80,000 2000 $80,000 0.76 $70,686 1.45 $102,495 $136,643

C-225 1 Lime Solids 

Feeder $3,900 1997 $3,900 $3,900 1.3 $5,070 $6,892

H-200 1 Hydrolyzate 

Cooler 1,988 2,288 1.15 $45,000 1997 $45,000 0.51 $48,340 2.1 $101,515 $137,997

H-201 2 1 Beer Column Feed 

Economizer 12,532 12,217 0.97 $132,800 1997 $398,400 0.68 $391,574 2.1 $822,304 $1,117,823

H-205 1 Pneumapress Vent 

Condensor 120 176 1.46 $15,385 2000 $15,385 0.68 $19,940 2.1 $41,874 $55,825

H-244 2 1 Waste Vapor 

Condensor 12,532 1,710 0.14 $132,800 1997 $398,400 0.68 $102,824 2.1 $215,931 $293,532

M-202 3 Prehydrolysis/Scre

w Feeder/Reactor 271,313 281,602 1.04 $2,454,982 2000 $7,364,947 0.6 $7,531,285 2.29 $17,246,643 $22,992,607

P-201 1 1 Sulfuric Acid 1,647 8,300 5.04 $4,800 1997 $9,600 0.79 $34,446 2.8 $96,450 $131,112
53

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

Pump

P-205 2 1 Pneumapress Feed 

Pump 50,299 42,588 0.85 $15,416 2000 $46,248 0.79 $40,551 3.34 $135,440 $180,563

P-209 1 1 Overlimed 

Hydrolyzate Pump 167,050 301,831 1.81 $10,700 1997 $21,400 0.79 $34,149 2.8 $95,617 $129,980

P-211 1 1 Primary Filtrate 

Pump 136,350 159,432 1.17 $32,549 2000 $65,098 0.79 $73,659 3.56 $262,225 $349,590

P-213 1 1 Wash Filtrate 

Pump 131,530 136,282 1.04 $49,803 2000 $99,606 0.79 $102,438 2.71 $277,607 $370,096

P-222 1 1 Filtered 

Hydrolyzate Pump 162,090 292,599 1.81 $10,800 1997 $21,600 0.79 $34,443 2.8 $96,441 $131,099

P-223 1 Lime Unloading 

Blower 547 6,117 11.18 $47,600 1998 $47,600 0.5 $159,176 1.4 $222,847 $300,600

P-224 2 1 Saccharification 

Feed Pump 358,810 416,118 1.16 $61,368 1998 $184,104 0.7 $204,225 2.8 $571,831 $771,348

P-239 1 1 Reacidified Liquor 

Pump 167,280 306,440 1.83 $10,800 1997 $21,600 0.79 $34,845 2.8 $97,567 $132,631

S-205 3 Pneumapress Filter 50,299 42,588 0.85 $1,575,000 2000 $4,725,000 0.6 $4,276,015 1.05 $4,489,816 $5,985,662

S-222 1 Hydroclone & 

Rotary Drum Filter 5,195 13,841 2.66 $165,000 1998 $165,000 0.39 $241,801 1.4 $338,521 $456,634

S-227 1 LimeDust Vent 

Baghouse 548 6,117 11.16 $32,200 1997 $32,200 1 $359,422 1.5 $539,133 $732,886

T-201 1 Sulfuric Acid Tank 1,647 8,300 5.04 $5,760 1996 $5,760 0.71 $18,160 1.4 $25,423 $34,995

T-203 1 Blowdown Tank 270,300 281,602 1.04 $64,100 1997 $64,100 0.93 $66,589 1.2 $79,907 $108,624

T-205 1 Hydrolyzate 

Mixing Tank 358,810 236,880 0.66 $44,800 1997 $44,800 0.71 $33,361 1.2 $40,033 $54,420

T-209 1 Overliming Tank 167,050 301,831 1.81 $71,000 1997 $71,000 0.71 $108,061 1.4 $151,286 $205,654

T-211 1 Primary Filtrate 

Tank 136,350 159,432 1.17 $36,000 2000 $36,000 0.71 $40,228 2.45 $98,558 $131,394

T-213 1 Wash Filtrate Tank 131,530 136,282 1.04 $18,000 2000 $18,000 0.71 $18,459 3.68 $67,930 $90,562

T-220 1 Lime Storage Bin 548 6,117 11.16 $69,200 1997 $69,200 0.46 $209,928 1.3 $272,906 $370,983

T-224 1 Reacidification 

Tank 167,280 306,440 1.83 $147,800 1997 $147,800 0.51 $201,258 1.2 $241,510 $328,304

T-232 1 Slurrying Tank 358,810 416,118 1.16 $44,800 1997 $44,800 0.71 $49,770 1.2 $59,724 $81,188
54

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

A200 Subtotal $14,464,84

8 $14,757,407 1.84 $27,102,397 $36,235,330

A-300 12 Ethanol Fermentor 

Agitator $19,676 1996 $236,112 $236,112 1.2 $283,334 $390,002

A-301 1 Seed Hold Tank 

Agitator 41,777 40,806 0.98 $12,551 1996 $12,551 0.51 $12,401 1.2 $14,882 $20,484

A-304 2 4th Seed Vessel 

Agitator 41,777 40,806 0.98 $11,700 1997 $23,400 0.51 $23,121 1.2 $27,745 $37,716

A-305 2 5th Seed Vessel 

Agitator 41,777 40,806 0.98 $10,340 1996 $20,680 0.51 $20,433 1.2 $24,520 $33,751

A-306 2 Beer Surge Tank 

Agitator 381,700 429,208 1.12 $48,700 1998 $97,400 0.51 $103,405 1.2 $124,086 $167,381

A-310 30 Saccharification 

Tank Agitator $19,676 1996 $590,280 $590,280 1.2 $708,336 $975,006

F-300 6 Ethanol Fermentor $493,391 1998 $2,960,346 $2,960,346 1.2 $3,552,415 $4,791,884

F-301 2 1st Seed 

Fermentor $14,700 1997 $29,400 $29,400 2.8 $82,320 $111,904

F-302 2 2nd Seed 

Fermentor $32,600 1997 $65,200 $65,200 2.8 $182,560 $248,168

F-303 2 3rd Seed 

Fermentor $81,100 1997 $162,200 $162,200 2.8 $454,160 $617,376

F-304 2 4th Seed 

Fermentor 41,777 40,806 0.98 $39,500 1997 $79,000 0.93 $77,291 1.2 $92,749 $126,081

F-305 2 5th Seed 

Fermentor 41,777 40,806 0.98 $147,245 1998 $294,490 0.51 $290,978 1.2 $349,174 $471,004

H-300 6 1 Fermentation 

Cooler 67,820 129,602 1.91 $4,000 1997 $28,000 0.78 $46,402 2.1 $97,444 $132,463

H-301 1 1 Hydrolyzate 

Heater 256 376 1.47 $22,400 2001 $44,800 0.68 $58,229 2.1 $122,282 $162,939

H-302 3 Saccharified Slurry 

Cooler 3,765 0 0.00 $25,409 1998 $76,227 0.78 $0 2.1 $0 $0

H-304 1 4th Seed 

Fermentor Coil 0.1380 0.2846 2.06 $3,300 1997 $3,300 0.83 $6,017 1.2 $7,220 $9,815

H-305 1 5th Seed 

Fermentor Coil 0.1380 0.2846 2.06 $18,800 1997 $18,800 0.98 $38,208 1.2 $45,849 $62,326

H-310 15 1 Saccharification 

Cooler 67,820 1,696 0.03 $4,000 1997 $64,000 0.78 $3,603 2.1 $7,566 $10,286
55

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

P-300 6 1

Fermentation 

Recirc/Transfer 

Pump

67,737 129,602 1.91 $8,000 1997 $56,000 0.79 $93,497 2.8 $261,792 $355,875

P-301 1 1 Seed Hold 

Transfer Pump 41,777 40,806 0.98 $22,194 1998 $44,388 0.7 $43,663 1.4 $61,128 $82,457

P-302 2 Seed Transfer 

Pump 41,777 40,806 0.98 $54,088 1998 $108,176 0.7 $106,409 1.4 $148,973 $200,951

P-306 1 1 Beer Transfer 

Pump 381,701 429,208 1.12 $17,300 1997 $34,600 0.79 $37,960 2.8 $106,287 $144,484

P-310 15 1

Saccharification 

Recirc/Transfer 

Pump

67,737 1,696 0.03 $8,000 1997 $128,000 0.79 $6,952 2.8 $19,465 $26,461

T-301 1 Seed Hold Tank 41,777 40,806 0.98 $161,593 1998 $161,593 0.51 $159,666 1.2 $191,599 $258,450

T-306 1 Beer Storage Tank 381,700 429,208 1.12 $237,700 1998 $237,700 0.71 $258,345 1.2 $310,014 $418,181

T-310 15 Saccharification 

Tank $493,391 1998 $7,400,865 $7,400,865 1.2 $8,881,038 $11,979,711

A300 Subtotal $12,977,508 $12,830,984 1.3 $16,156,940 $21,835,156

A400 Subtotal $0 $0 $0 $0

A-530 1 Recycled Water 

Tank Agitator 179,446 263,502 1.47 $5,963 1998 $5,963 0.51 $7,254 1.3 $9,430 $12,720

C-501 1 Lignin Wet Cake 

Screw 99,199 47,373 0.48 $31,700 1997 $31,700 0.78 $17,811 1.4 $24,936 $33,897

D-501 1 Beer Column 17 20 1.17 $478,100 1998 $478,100 0.68 $531,355 2.1 $1,115,846 $1,505,174

D-502 1 Rectification 

Column 56,477 62,095 1.10 $525,800 1996 $525,800 0.68 $560,821 2.1 $1,177,724 $1,621,105

E-501 2 1st Effect 

Evaporation 22,278 32,408 1.45 $544,595 1996 $1,089,190 0.68 $1,405,364 2.1 $2,951,265 $4,062,339

E-502 1 2nd Effect 

Evaporation 22,278 32,408 1.45 $435,650 1996 $435,650 0.68 $562,112 2.1 $1,180,436 $1,624,839

E-503 2 3rd Effect 

Evaporation 22,278 32,408 1.45 $435,650 1996 $871,300 0.68 $1,124,224 2.1 $2,360,871 $3,249,677

H-501 1 1 Beer Column 

Reboiler

-

28.3092 -32 1.14 $158,374 1996 $316,748 0.68 $346,704 2.1 $728,079 $1,002,181
56

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

H-502 1 Rectification 

Column Reboiler -3.5547 -4 0.99 $29,600 1997 $29,600 0.68 $29,382 2.1 $61,702 $83,876

H-504 1 Beer Column 

Condenser 1.0002 0 0.45 $29,544 1996 $29,544 0.68 $17,260 2.1 $36,246 $49,892

H-505 1 Start-up Rect. 

Column Condenser 17.6595 19 1.06 $86,174 1996 $86,174 0.68 $89,788 2.1 $188,555 $259,542

H-512 1 1 Beer Column Feed 

Interchanger 909 914 1.01 $19,040 1996 $38,080 0.68 $38,214 2.1 $80,250 $110,462

H-517 1 1 Evaporator 

Condenser 24 33 1.35 $121,576 1996 $243,152 0.68 $297,600 2.1 $624,959 $860,240

M-503 1 Molecular Sieve (9 

pieces) 20,491 19,052 0.93 $2,700,000 1998 $2,700,000 0.7 $2,565,867 1 $2,565,867 $3,461,120

P-501 1 1 Beer Column 

Bottoms Pump 5,053 6,003 1.19 $42,300 1997 $84,600 0.79 $96,938 2.8 $271,425 $368,970

P-503 1 1 Beer Column 

Reflux Pump 1.0002 0.4537 0.45 $1,357 1998 $2,714 0.79 $1,454 2.8 $4,070 $5,490

P-504 1 1

Rectification 

Column Bottoms 

Pump

31,507 36,584 1.16 $4,916 1998 $9,832 0.79 $11,064 2.8 $30,978 $41,787

P-505 1 1

Rectification 

Column Reflux 

Pump

18 19 1.06 $4,782 1998 $9,564 0.79 $10,030 2.8 $28,084 $37,883

P-511 2 1 1st Effect Pump 278,645 311,140 1.12 $19,700 1997 $59,100 0.79 $64,481 2.8 $180,547 $245,432

P-512 1 1 2nd Effect Pump 91,111 137,819 1.51 $13,900 1997 $27,800 0.79 $38,551 2.8 $107,944 $146,736

P-513 2 1 3rd Effect Pump 48,001 78,409 1.63 $8,000 1997 $24,000 0.79 $35,365 2.8 $99,022 $134,608

P-514 1 1 Evaporator 

Condensate Pump 140,220 181,051 1.29 $12,300 1997 $24,600 0.79 $30,104 2.8 $84,290 $114,582

P-515 1 Scrubber Bottoms 

Pump 15,377 18,179 1.18 $2,793 1998 $2,793 0.79 $3,188 2.8 $8,926 $12,040

P-530 1 1 Recycled Water 

Pump 179,446 263,502 1.47 $10,600 1997 $21,200 0.79 $28,718 2.8 $80,409 $109,306

S-505 4 Pneumapress Filter 26,601 18,737 0.70 $1,418,750 2000 $5,675,000 0.6 $4,598,800 1.04 $4,782,752 $6,376,193

T-503 1 Beer Column 

Relfux Drum 1.0002 0 0.45 $11,900 1997 $11,900 0.93 $5,706 2.1 $11,982 $16,288

T-505 1 Rectification 

Column Reflux 17.6627 19 1.06 $45,600 1997 $45,600 0.72 $47,621 2.1 $100,005 $135,944
57

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

Drum

T-512 1 Vent Scrubber 18,523 18,916 1.02 $99,000 1998 $99,000 0.78 $100,635 2.1 $211,335 $285,071

T-514 1 Evaporator 

Condensate Drum 164,760 181,051 1.10 $37,200 1998 $37,200 0.93 $40,609 2.1 $85,279 $115,034

T-530 1 Recycled Water 

Tank 179,446 263,502 1.47 $14,515 1998 $14,515 0.745 $19,325 1.4 $27,055 $36,495

A500 Subtotal $13,030,419 $12,726,345 1.51 $19,220,269 $26,118,926

A-602 1 Equalization Basin 

Agitator 188,129 85,599 0.46 $28,400 1997 $28,400 0.51 $19,007 1.2 $22,808 $31,005

A-606 1 Anaerobic Agitator 810,250 401,778 0.50 $30,300 1997 $30,300 0.51 $21,188 1.2 $25,425 $34,562

A-608 16 Aerobic Lagoon 

Agitator 812 108 0.13 $31,250 1998 $500,000 0.51 $178,402 1.4 $249,763 $336,908

C-614 1 Aerobic Sludge 

Screw 978 112 0.11 $5,700 1997 $5,700 0.78 $1,051 1.4 $1,471 $2,000

H-602 1

Anaerobic 

Digestor Feed 

Cooler

7,627 3,155 0.41 $128,600 1997 $128,600 0.74 $66,913 2.1 $140,518 $191,017

M-604 1 Nutrient Feed 

System $31,400 1998 $31,400 $31,400 2.58 $81,012 $109,278

M-606 1 Biogas Emergency 

Flare 2,572 292 0.11 $20,739 1998 $20,739 0.6 $5,620 1.68 $9,442 $12,737

M-612 1 Filter Precoat 

System $3,000 1998 $3,000 $3,000 1.4 $4,200 $5,665

P-602 1 1 Anaerobic Reactor 

Feed Pump 188,129 85,599 0.46 $11,400 1997 $22,800 0.79 $12,240 2.8 $34,271 $46,587

P-606 1 1 Aerobic Digestor 

Feed Pump 185,782 85,335 0.46 $10,700 1997 $21,400 0.79 $11,574 2.8 $32,408 $44,055

P-608 1 Aerobic Sludge 

Recycle Pump 5,862 679 0.12 $11,100 1997 $11,100 0.79 $2,022 1.4 $2,831 $3,848

P-610 1 Aerobic Sludge 

Pump 5,862 679 0.12 $11,100 1997 $11,100 0.79 $2,022 1.4 $2,831 $3,848

P-611 1 1 Aerobic Digestion 

Outlet Pump 187,827 85,560 0.46 $10,700 1997 $21,400 0.79 $11,498 2.8 $32,196 $43,766

P-614 1 1 Sludge Filtrate 

Recycle Pump 4,885 567 0.12 $6,100 1997 $12,200 0.79 $2,226 2.8 $6,234 $8,474
58

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

P-616 1 1 Treated Water 

Pump 181,965 84,881 0.47 $10,600 1997 $21,200 0.79 $11,607 2.8 $32,498 $44,178

S-600 1 Bar Screen 188,129 85,599 0.46 $117,818 1991 $117,818 0.3 $93,028 1.2 $111,634 $162,338

S-614 1 Belt Filter Press 438 58 0.13 $650,223 1998 $650,223 0.72 $151,646 1.8 $272,963 $368,203

T-602 1 Equalization Basin 188,129 85,599 0.46 $350,800 1998 $350,800 0.51 $234,772 1.42 $333,376 $449,694

T-606 1 Anaerobic 

Digestor 810,250 401,778 0.50 $881,081 1998 $881,081 0.51 $616,103 1.04 $640,747 $864,309

T-608 1 Aerobic Digestor 19,506,756 8,960,675 0.46 $635,173 1998 $635,173 1 $291,775 1 $291,775 $393,578

T-610 1 Clarifier 185,782 85,335 0.46 $174,385 1998 $174,385 0.51 $117,272 1.96 $229,852 $310,050

A600 Subtotal $3,678,819 $1,884,366 1.36 $2,558,255 $3,466,097

A-701 1 Denaturant In-line 

Mixer 19,436 19,838 1.02 $1,900 1997 $1,900 0.48 $1,919 1 $1,919 $2,608

A-720 1 CSL Storage Tank 

Agitator 41,777 4,488 0.11 $12,551 1996 $12,551 0.51 $4,023 1.2 $4,828 $6,645

A-760 1 CSL/DAP Day 

Tank Agitator 1,400 4,488 3.21 $12,795 2001 $12,795 0.51 $23,177 1.2 $27,813 $37,060

C-755 1 DAP Solids Feeder $3,900 1997 $3,900 $3,900 1.3 $5,070 $6,892

P-701 2 1 Ethanol Product 

Pump 18,549 19,052 1.03 $7,500 1997 $22,500 0.79 $22,981 2.8 $64,347 $87,472

P-703 1 1 Sulfuric Acid 

Pump 1,647 8,300 5.04 $8,000 1997 $16,000 0.79 $57,411 2.8 $160,750 $218,520

P-704 1 1 Firewater Pump 6,823 7,008 1.03 $18,400 1997 $36,800 0.79 $37,585 2.8 $105,238 $143,059

P-710 1 1 Gasoline Pump 887 786 0.89 $4,500 1997 $9,000 0.79 $8,181 2.8 $22,907 $31,139

P-720 1 1 CSL Pump 2,039 4,488 2.20 $8,800 1997 $17,600 0.79 $32,824 2.8 $91,908 $124,937

P-750 1 1 Cellulase Pump 6,823 8,712 1.28 $18,400 1997 $36,800 0.79 $44,637 2.8 $124,983 $169,900

P-755 1 DAP Unloading 

Blower 154 30 0.19 $47,600 1998 $47,600 0.5 $20,922 1.4 $29,291 $39,511

P-760 1 1 CSL/DAP Pump 2,039 4,488 2.20 $8,800 1997 $17,600 0.79 $32,824 2.8 $91,908 $124,937

S-755 1 DAP Vent 

Baghouse 548 30 0.05 $32,200 1997 $32,200 1 $1,748 1.5 $2,622 $3,565
59

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

T-701 2 Ethanol Product 

Storage Tank 18,549 19,052 1.03 $165,800 1997 $331,600 0.51 $336,159 1.4 $470,623 $639,755

T-703 1 Sulfuric Acid 

Storage Tank 1,647 8,300 5.04 $42,500 1997 $42,500 0.51 $96,961 1.2 $116,353 $158,168

T-704 1 Firewater Storage 

Tank 6,822 7,008 1.03 $166,100 1997 $166,100 0.51 $168,392 1.4 $235,748 $320,472

T-709 1 Propane Storage 

Tank 15 16 1.04 $24,834 2001 $24,834 0.72 $25,580 1.4 $35,812 $47,719

T-710 1 Gasoline Storage 

Tank 887 786 0.89 $43,500 1997 $43,500 0.51 $40,902 1.4 $57,262 $77,841

T-720 1 CSL Storage Tank 2,039 4,488 2.20 $88,100 1997 $88,100 0.79 $164,307 1.4 $230,030 $312,699

T-750 2 Cellulase Storage 

Tank 9,234 8,712 0.94 $125,900 2001 $251,800 0.79 $240,483 1.4 $336,676 $448,617

T-755 1 DAP Storage Bin 154 30 0.19 $33,384 2001 $33,384 0.44 $16,195 1.3 $21,054 $28,054

T-760 1 CSL/DAP Day 

Tank 1,400 4,488 3.21 $30,084 2001 $30,084 0.79 $75,512 1.4 $105,717 $140,866

A700 Subtotal $1,279,148 $1,456,624 1.6 $2,342,860 $3,170,437

H-801 1

Burner 

Combustion Air 

Preheater

24.0337 13 0.56 $1,049,900 1997 $1,049,900 0.6 $739,305 1.5 $1,108,957 $1,507,493

H-811 1 BFW Preheater 415 226 0.54 $58,400 1997 $58,400 0.68 $38,647 2.1 $81,159 $110,325

M-803 1

Fluidized Bed 

Combustion 

Reactor

341,270 209,403 0.61 24,900,000 1998 $24,900,00

0 0.75 $17,262,891 1.3 $22,441,758 $30,271,886

M-804 1 Combustion Gas 

Baghouse 652,517 450,947 0.69 $2,536,300 1998 $2,536,300 0.58 $2,047,058 1.5 $3,070,587 $4,141,941

M-811 1 Turbine/Generator 281,179 209,403 0.74 10000000 1998 $10,000,00

0 0.71 $8,111,862 1.5 $12,167,793 $16,413,244

M-820 1 Hot Process Water 

Softener System 225,889 192,701 0.85 $1,381,300 1999 $1,381,300 0.82 $1,212,546 1.3 $1,576,310 $2,120,311

M-830 1 Hydrazine 

Addition Pkg. 229,386 215,880 0.94 $19,000 1994 $19,000 0.6 $18,321 1 $18,321 $26,150

M-832 1 Ammonia 

Addition Pkg 229,386 215,880 0.94 $19,000 1994 $19,000 0.6 $18,321 1 $18,321 $26,150
60

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

M-834 1 Phosphate 

Addition Pkg. 229,386 215,880 0.94 $19,000 1994 $19,000 0.6 $18,321 1 $18,321 $26,150

P-804 2 Condensate Pump 61,471 136,035 2.21 $7,100 1997 $14,200 0.79 $26,596 2.8 $74,470 $101,233

P-811 2 Turbine 

Condensate Pump 39,524 59,037 1.49 $7,800 1997 $15,600 0.79 $21,419 2.8 $59,972 $81,525

P-824 2 Deaerator Feed 

Pump 293,605 192,701 0.66 $9,500 1997 $19,000 0.79 $13,623 2.8 $38,145 $51,853

P-826 5 BFW Pump 564,626 215,880 0.38 $52,501 1998 $262,505 0.79 $122,821 2.8 $343,900 $463,890

P-828 2 Blowdown Pump 6,613 6,476 0.98 $5,100 1997 $10,200 0.79 $10,033 2.8 $28,093 $38,189

P-830 1 Hydrazine 

Transfer Pump 229,386 215,880 0.94 $5,500 1997 $5,500 0.79 $5,243 2.8 $14,679 $19,955

T-804 1 Condensate 

Collection Tank 228,862 136,035 0.59 $7,100 1997 $7,100 0.71 $4,907 1.4 $6,870 $9,339

T-824 1 Condensate Surge 

Drum 222,360 192,701 0.87 $49,600 1997 $49,600 0.72 $44,742 1.7 $76,062 $103,397

T-826 1 Deaerator 266,213 215,880 0.81 $165,000 1998 $165,000 0.72 $141,890 2.8 $397,292 $535,911

T-828 1 Blowdown Flash 

Drum 6,563 6,476 0.99 $9,200 1997 $9,200 0.72 $9,112 2.8 $25,515 $34,684

T-830 1 Hydrazine Drum 229,386 215,880 0.94 $12,400 1997 $12,400 0.93 $11,720 1.7 $19,923 $27,083

A800 Subtotal $40,553,205 $29,879,378 1.4 $41,586,447 $56,110,709

M-902 1 Cooling Tower 

System

147.960

0 105 0.71 $1,659,000 1998 $1,659,000 0.78 $1,265,253 1.2 $1,518,304 $2,048,054

M-904 2 1 Plant Air 

Compressor 98,040 111,111 1.13 $278,200 2000 $834,600 0.34 $870,881 1.3 $1,132,145 $1,509,336

M-910 1 CIP System 63 63 1.00 $95,000 1995 $95,000 0.6 $95,036 1.2 $114,043 $157,225

P-902 1 1 Cooling Water 

Pump

182900

00

12,447,

929 0.68 $332,300 1997 $664,600 0.79 $490,386 2.8 $1,373,081 $1,866,538

P-912 1 1 Make-up Water 

Pump 244,160 211,894 0.87 $10,800 1997 $21,600 0.79 $19,312 2.8 $54,073 $73,506

P-914 2 1 Process Water 

Circulating Pump 352,710 254,604 0.72 $11,100 1997 $33,300 0.79 $25,741 2.8 $72,073 $97,975

S-904 1 1 Instrument Air 

Dryer 159,950 111,111 0.69 $15,498 1999 $30,996 0.6 $24,910 1.3 $32,383 $43,558
61

Equipment IDNumber 

Required

Spares 

Nos Equipment Name

Scaling 

Stream 

Flow 

(Kg/hr)

New 

Stream 

Flow

Size 

Ratio

Original 

Equip Cost 

(per unit)

Base 

Year

Total 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year

Scaling 

Exponent

Scaled Cost in 

Base Year

Installation 

Factor

Installed Cost 

in Base Year

Installed Cost 

in 2007$

T-902 3 Prehydrolysis 

Filter Air Receiver 5,259 4,687 0.89 $17,000 2000 $51,000 0.72 $46,942 1.2 $56,331 $75,098

T-904 1 Plant Air Receiver 159,950 111,111 0.69 $13,000 1997 $13,000 0.72 $10,000 1.3 $13,001 $17,673

T-905 4 Product Recovery 

Filter Air Receiver 5,700 5,921 1.04 $17,000 2000 $68,000 0.72 $69,887 1.2 $83,865 $111,806

T-914 1 Process Water 

Tank 352,710 254,604 0.72 $195,500 1997 $195,500 0.51 $165,560 1.4 $231,784 $315,082

A900 Subtotal $3,666,596 $3,083,909 1.5 $4,681,083 $6,315,850

Equipment Cost $93,786,061 $81,109,704 1.50 $121,788,091 $164,104,477
62

Appendix F - Process Operating Summaries

Table F-1. Operating Summary for Dilute Acid Pretreatment Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0190 Enzyme Loading (mg/g cell) 35.1 Total Solids (wt%) 20.1%

Acid Loading (g acid/g dry biomass) 0.0445 Total Solids (wt%) 20.0% Insoluble Solids (wt%) 4.3%

Total Solids (wt%) 0.2956 Insoluble Solids (wt%) 10.0% Temperature (°C) 32

Temperature (°C) 190 Temperature (°C) 32 Pressure (atm) 1.0

Pressure (atm) 11.5 Pressure (atm) 1.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 5 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.95

 Cellulose to Glucolig 0.003 Cellulose to Glucolig 0 Glucose to Zymo 0.02

 Cellulose to Cellobiose 0 Cellulose to Cellobiose 0 Glucose to Glycerol 0.004

 Cellulose to Glucose 0.099 Cellulose to Glucose 0.9109 Glucose to Succinic Acid 0.006

 Cellulose to HMF 0.003 Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.015

 Xylan to Oligomer 0.21 Glucolig to Glucose 0 Glucose to Lactic Acid 0.002

 Xylan to Xylose 0.6 Cellobiose to Glucose 0 Xylose to Ethanol 0.756

 Xylan to Furfural 0.11 Xylan to Oligomer 0 Xylose to Zymo 0.019

 Xylan to Tar 0 Xylan to Xylose 0.5713 Xylose to Glycerol 0.003

 Mannan to Oligomer 0.21 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0.046

 Mannan to Mannose 0.6 Xylan to Tar 0 Xylose to Succinic Acid 0.009

 Mannan to HMF 0.08 Arabinan to Oligomer 0.5713 Xylose to Acetic Acid 0.014

 Galactan to Oligomer 0.21 Arabinan to Arabinose 0 Xylose to Lactic Acid 0.002

 Galactan to Galactose 0.6 Galactan to Oligomer 0 Arabinose to Ethanol 0

 Galactan to HMF 0.08 Galactan to Galactose 0.5713 Arabinose to Zymo 0

 Arabinan to Oligomer 0.21 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Arabinose 0.6 Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Arabinan to Furfural 0.08 Mannan to Mannose 0.5713 Arabinose to Acetic Acid 0

 Arabinan to Tar 0 Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Acetate to Oligomer 0 Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 Acetate to Acetic Acid 1 Soluble Sugars From PT 22,989 Galactose to Zymo 0

 Furfural to Tar 1 Other Soluble Solids From PT 18,543 Galactose to Glycerol 0

 HMF to Tar 1 Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.1 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0
63

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

S/L Separation Soluble Sugars in Produced Cellulase 0 Galactose to Lactic Acid 0

Water/Hydrolyzate Ratio (kg/kg) 0.58 Other Sol Solids in Produced Cellulase 0 Mannose to Ethanol 0

Water to S/L Separator (kg/hr) 131,829 Mannose to Zymo 0

Dilution Water (kg/hr) 52,436 Soluble Sugars From Sacc 44,457 Mannose to Glycerol 0

Conditioning Other Soluble Solids From Sacc 16,688 Mannose to Succinic Acid 0

Ca(OH)2 to Gypsum 0.145 Soluble Sugars From Seed Train 880 Mannose to Acetic Acid 0

Sugar Conversion to TAR: Other Soluble Solids From Seed Train 1,955 Mannose to Lactic Acid 0

 Xylose 0.13 Soluble Sugars From DAP 0

 Arabinose 0.2 Other Soluble Solids From DAP 0 Contamination Loss 7.0%

 Glucose 0.12 Soluble Sugars From CSL 0

 Galactose 0.28 Other Soluble Solids From CSL 2,095 Ethanol Out of Fermenters (wt%) 4.7%

 Mannose 0

 Cellobiose 0.36
64

Table F-2. Operating Summary for Dilute Acid Pretreatment (Pilot) Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0190 Enzyme Loading (mg/g cell) 35.1 Total Solids (wt%) 20.1%

Acid Loading (g acid/g dry biomass) 0.0445 Total Solids (wt%) 20.0% Insoluble Solids (wt%) 4.3%

Total Solids (wt%) 0.2956 Insoluble Solids (wt%) 10.0% Temperature (°C) 32

Temperature (°C) 190 Temperature (°C) 32 Pressure (atm) 1.0

Pressure (atm) 11.5 Pressure (atm) 1.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 5 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.95

 Cellulose to Glucolig 0.003 Cellulose to Glucolig 0 Glucose to Zymo 0.02

 Cellulose to Cellobiose 0 Cellulose to Cellobiose 0 Glucose to Glycerol 0.004

 Cellulose to Glucose 0.099 Cellulose to Glucose 0.9109 Glucose to Succinic Acid 0.006

 Cellulose to HMF 0.003 Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.015

 Xylan to Oligomer 0.21 Glucolig to Glucose 0 Glucose to Lactic Acid 0.002

 Xylan to Xylose 0.6 Cellobiose to Glucose 0 Xylose to Ethanol 0.756

 Xylan to Furfural 0.11 Xylan to Oligomer 0 Xylose to Zymo 0.019

 Xylan to Tar 0 Xylan to Xylose 0.5713 Xylose to Glycerol 0.003

 Mannan to Oligomer 0.21 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0.046

 Mannan to Mannose 0.6 Xylan to Tar 0 Xylose to Succinic Acid 0.009

 Mannan to HMF 0.08 Arabinan to Oligomer 0.5713 Xylose to Acetic Acid 0.014

 Galactan to Oligomer 0.21 Arabinan to Arabinose 0 Xylose to Lactic Acid 0.002

 Galactan to Galactose 0.6 Galactan to Oligomer 0 Arabinose to Ethanol 0

 Galactan to HMF 0.08 Galactan to Galactose 0.5713 Arabinose to Zymo 0

 Arabinan to Oligomer 0.21 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Arabinose 0.6 Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Arabinan to Furfural 0.08 Mannan to Mannose 0.5713 Arabinose to Acetic Acid 0

 Arabinan to Tar 0 Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Acetate to Oligomer 0 Galactose to Ethanol 0

 Acetate to Acetic Acid 1 Sugar & Solids Flow Rates (kg/hr) Galactose to Zymo 0

 Furfural to Tar 1 Soluble Sugars From PT 22,989 Galactose to Glycerol 0

 HMF to Tar 1 Other Soluble Solids From PT 18,543 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.1 Soluble Sugars in Purchased Cellulase 0 Galactose to Acetic Acid 0

Other Sol Solids in Purchased Cellulase 0 Galactose to Lactic Acid 0

Soluble Sugars in Produced Cellulase 0 Mannose to Ethanol 0

Other Sol Solids in Produced Cellulase 0 Mannose to Zymo 0
65

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

S/L Separation Soluble Sugars From Sacc 44,457 Mannose to Glycerol 0

Water/Hydrolyzate Ratio (kg/kg) 0.58 Other Soluble Solids From Sacc 16,688 Mannose to Succinic Acid 0

Water to S/L Separator (kg/hr) 131,829 Soluble Sugars From Seed Train 880 Mannose to Acetic Acid 0

Dilution Water (kg/hr) 52,436 Other Soluble Solids From Seed Train 1,955 Mannose to Lactic Acid 0

Conditioning Soluble Sugars From DAP 0

Ca(OH)2 to Gypsum 0.145 Other Soluble Solids From DAP 0 Contamination Loss 7.0%

Sugar Conversion to TAR: Soluble Sugars From CSL 0

 Xylose 0.13 Other Soluble Solids From CSL 2,095 Ethanol Out of Fermenters (wt%) 4.7%

 Arabinose 0.2

 Glucose 0.12

 Galactose 0.28

 Mannose 0

 Cellobiose 0.36
66

Table F-3. Operating Summary for Two-Stage Dilute Acid Pretreatment Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0190 Enzyme Loading (mg/g cell) N/A Total Solids (wt%) 19.7%

Acid Loading (g acid/g dry biomass) 0.0443 Total Solids (wt%) 34.2% Insoluble Solids (wt%) 7.1%

Total Solids (wt%) 0.2956 Insoluble Solids (wt%) 28.6% Temperature (°C) 32

Temperature (°C) 190 Temperature (°C) 144.0873 Pressure (atm) 1.0

Pressure (atm) 11.5 Pressure (atm) 4.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 0 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.95

 Cellulose to Glucolig 0 Cellulose to Glucolig N/A Glucose to Zymo 0.02

 Cellulose to Cellobiose 0 Cellulose to Cellobiose N/A Glucose to Glycerol 0.004

 Cellulose to Glucose 0.0626 Cellulose to Glucose N/A Glucose to Succinic Acid 0.006

 Cellulose to HMF 0 Glucolig to Cellobiose N/A Glucose to Acetic Acid 0.015

 Xylan to Oligomer 0.0265 Glucolig to Glucose N/A Glucose to Lactic Acid 0.002

 Xylan to Xylose 0.8249 Cellobiose to Glucose N/A Xylose to Ethanol 0.756

 Xylan to Furfural 0 Xylan to Oligomer N/A Xylose to Zymo 0.019

 Xylan to Tar 0 Xylan to Xylose N/A Xylose to Glycerol 0.003

 Mannan to Oligomer 0.0265 Xylose Oligomer to Xylose N/A Xylose to Xylitol 0.046

 Mannan to Mannose 0.8249 Xylan to Tar N/A Xylose to Succinic Acid 0.009

 Mannan to HMF 0 Arabinan to Oligomer N/A Xylose to Acetic Acid 0.014

 Galactan to Oligomer 0.0265 Arabinan to Arabinose N/A Xylose to Lactic Acid 0.002

 Galactan to Galactose 0.8249 Galactan to Oligomer N/A Arabinose to Ethanol 0

 Galactan to HMF 0 Galactan to Galactose N/A Arabinose to Zymo 0

 Arabinan to Oligomer 0.0265 Galactose Oligomer to Galactose N/A Arabinose to Glycerol 0

 Arabinan to Arabinose 0.8249 Mannan to Oligomer N/A Arabinose to Succinic Acid 0

 Arabinan to Furfural 0 Mannan to Mannose N/A Arabinose to Acetic Acid 0

 Arabinan to Tar 0 Mannose Oligomer to Mannose N/A Arabinose to Lactic Acid 0

 Acetate to Oligomer 0 Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 Acetate to Acetic Acid 1 Soluble Sugars From PT 0 Galactose to Zymo 0

 Furfural to Tar 1 Other Soluble Solids From PT 13,600 Galactose to Glycerol 0

 HMF to Tar 1 Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.1 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0

S/L Separation Soluble Sugars in Produced Cellulase 0 Galactose to Lactic Acid 0

Water/Hydrolyzate Ratio (kg/kg) 0.58 Other Sol Solids in Produced Cellulase 0 Mannose to Ethanol 0

Water to S/L Separator (kg/hr) 137,458 Mannose to Zymo 0
67

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

Conditioning Soluble Sugars From Sacc 33,169 Mannose to Glycerol 0

Ca(OH)2 to Gypsum 0.183 Other Soluble Solids From Sacc 16,462 Mannose to Succinic Acid 0

Sugar Conversion to TAR: Soluble Sugars From Seed Train 881 Mannose to Acetic Acid 0

 Xylose 0.13 Other Soluble Solids From Seed Train 1,934 Mannose to Lactic Acid 0

 Arabinose 0.2 Soluble Sugars From DAP 0

 Glucose 0.12 Other Soluble Solids From DAP 0 Contamination Loss 7.0%

 Galactose 0.28 Soluble Sugars From CSL 0

 Mannose 0 Other Soluble Solids From CSL 2,159 Ethanol Out of Fermenters (wt%) 3.1%

 Cellobiose 0.36
68

Table F-4. Operating Summary for Hot Water Pretreatment Processes

Pretreatment Saccharification Fermentation

Total Solids (wt%) 17.2% Enzyme Loading (mg/g cell) 33.1 Total Solids (wt%) 16.5%

Total Insoluble Solids (wt%) 12.9% Total Solids (wt%) 16.0% Insoluble Solids (wt%) 3.9%

Temperature (°C) 190.0 Insoluble Solids (wt%) 9.2% Temperature (°C) 32

Pressure (atm) 12.3 Temperature (°C) 32 Pressure (atm) 1.0

Conversions: Pressure (atm) 1.0 Residence Time (days) 2

 Cellulose to Glucolig 0.053 Residence Time (days) 5 Conversions:

 Cellulose to Cellobiose 0 Conversions: Glucose to Ethanol 0.95

 Cellulose to Glucose 0.0032 Cellulose to Glucolig 0 Glucose to Zymo 0.02

 Cellulose to HMF 0 Cellulose to Cellobiose 0 Glucose to Glycerol 0.004

 Xylan to Oligomer 0.554 Cellulose to Glucose 0.8997 Glucose to Succinic Acid 0.006

 Xylan to Xylose 0.0239 Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.015

 Xylan to Furfural 0 Glucolig to Glucose 0 Glucose to Lactic Acid 0.002

 Xylan to Tar 0 Cellobiose to Glucose 0 Xylose to Ethanol 0.756

 Mannan to Oligomer 0.554 Xylan to Oligomer 0 Xylose to Zymo 0.019

 Mannan to Mannose 0.0239 Xylan to Xylose 0.5661 Xylose to Glycerol 0.003

 Mannan to HMF 0 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0.046

 Galactan to Oligomer 0.554 Xylan to Tar 0 Xylose to Succinic Acid 0.009

 Galactan to Galactose 0.0239 Arabinan to Oligomer 0.5661 Xylose to Acetic Acid 0.014

 Galactan to HMF 0 Arabinan to Arabinose 0 Xylose to Lactic Acid 0.002

 Arabinan to Oligomer 0.554 Galactan to Oligomer 0 Arabinose to Ethanol 0

 Arabinan to Arabinose 0.0239 Galactan to Galactose 0.5661 Arabinose to Zymo 0

 Arabinan to Furfural 0 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Tar 0 Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Acetate to Oligomer 0 Mannan to Mannose 0.5661 Arabinose to Acetic Acid 0

 Acetate to Acetic Acid 1 Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Furfural to Tar 1 Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 HMF to Tar 1 Soluble Sugars From PT 19,745 Galactose to Zymo 0

 Lignin to Soluble Lignin 0.05 Other Soluble Solids From PT 18,174 Galactose to Glycerol 0

 Glucose to HMF 0.5 Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Mannose to HMF 0.5 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0

 Galactose to HMF 0.5 Soluble Sugars in Produced Cellulase 0 Galactose to Lactic Acid 0

 Xylose to Furfural 0.5 Other Sol Solids in Produced Cellulase 0 Mannose to Ethanol 0

 Arabinose to Furfural 0.5 Mannose to Zymo 0
69

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

Soluble Sugars From Sacc 47,147 Mannose to Glycerol 0

S/L Separation Other Soluble Solids From Sacc 16,357 Mannose to Succinic Acid 0

Water/Hydrolyzate Ratio (kg/kg) #N/A Soluble Sugars From Seed Train 2,176 Mannose to Acetic Acid 0

Water to S/L Separator (kg/hr) #N/A Other Soluble Solids From Seed Train 1,954 Mannose to Lactic Acid 0

Dilution Water (kg/hr) #N/A Soluble Sugars From DAP 0

Conditioning Other Soluble Solids From DAP 0 Contamination Loss 7.0%

Ca(OH)2 to Gypsum 0.000 Soluble Sugars From CSL 0

Sugar Conversion to TAR: Other Soluble Solids From CSL 2,792 Ethanol Out of Fermenters (wt%) 2.7%

 Xylose 0

 Arabinose 0

 Glucose 0

 Galactose 0

 Mannose 0

 Cellobiose 0
70

Table F-5. Operating Summary for AFEX Pretreatment Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0000 Enzyme Loading (mg/g cell) 31.3 Total Solids (wt%) 20.5%

Acid Loading (g acid/g dry biomass) 0.0000 Total Solids (wt%) 20.0% Insoluble Solids (wt%) 3.6%

Total Solids (wt%) 0.4808 Insoluble Solids (wt%) 11.6% Temperature (°C) 32

Temperature (°C) 108.1302 Temperature (°C) 32 Pressure (atm) 1.0

Pressure (atm) 18.7 Pressure (atm) 1.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 5 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.95

 Cellulose to Glucolig Cellulose to Glucolig 0 Glucose to Zymo 0.02

 Cellulose to Cellobiose Cellulose to Cellobiose 0 Glucose to Glycerol 0.004

 Cellulose to Glucose Cellulose to Glucose 0.959 Glucose to Succinic Acid 0.006

 Cellulose to HMF Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.015

 Xylan to Oligomer 0.5 Glucolig to Glucose 0 Glucose to Lactic Acid 0.002

 Xylan to Xylose Cellobiose to Glucose 1 Xylose to Ethanol 0.756

 Xylan to Furfural Xylan to Oligomer 0 Xylose to Zymo 0.019

 Xylan to Tar Xylan to Xylose 0.777 Xylose to Glycerol 0.003

 Mannan to Oligomer 0.5 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0.046

 Mannan to Mannose Xylan to Tar 0 Xylose to Succinic Acid 0.009

 Mannan to HMF Arabinan to Oligomer 0.777 Xylose to Acetic Acid 0.014

 Galactan to Oligomer 0.5 Arabinan to Arabinose 0 Xylose to Lactic Acid 0.002

Galactan to Galactose Galactan to Oligomer 0 Arabinose to Ethanol 0

 Galactan to HMF Galactan to Galactose 0.777 Arabinose to Zymo 0

 Arabinan to Oligomer 0.5 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Arabinose Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Arabinan to Furfural Mannan to Mannose 0.777 Arabinose to Acetic Acid 0

 Arabinan to Tar Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Acetate to Oligomer Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 Acetate to Acetic Acid Soluble Sugars From PT 15,980 Galactose to Zymo 0

 Furfural to Tar Other Soluble Solids From PT 21,338 Galactose to Glycerol 0

 HMF to Tar Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.33 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0

S/L Separation Soluble Sugars in Produced Cellulase 0 Galactose to Lactic Acid 0

Water/Hydrolyzate Ratio (kg/kg) N/A Other Sol Solids in Produced Cellulase 0 Mannose to Ethanol 0
71

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

Water to S/L Separator (kg/hr) N/A Mannose to Zymo 0

Dilution Water (kg/hr) N/A Soluble Sugars From Sacc 50,400 Mannose to Glycerol 0

Conditioning Other Soluble Solids From Sacc 19,205 Mannose to Succinic Acid 0

Ca(OH)2 to Gypsum #N/A Soluble Sugars From Seed Train 1,962 Mannose to Acetic Acid 0

Sugar Conversion to TAR: Other Soluble Solids From Seed Train 2,242 Mannose to Lactic Acid 0

 Xylose #N/A Soluble Sugars From DAP 0

 Arabinose #N/A Other Soluble Solids From DAP 0 Contamination Loss 7.0%

 Glucose #N/A Soluble Sugars From CSL 0

 Galactose #N/A Other Soluble Solids From CSL 2,242 Ethanol Out of Fermenters (wt%) 3.9%

 Mannose #N/A

 Cellobiose #N/A
72

Table F-6. Operating Summary for Separate C5 & C6 Fermentation Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0190 Enzyme Loading (mg/g cell) 33.4 Total Solids (wt%) 17.1%

Acid Loading (g acid/g dry biomass) 0.0443 Total Solids (wt%) 16.1% Insoluble Solids (wt%) 3.9%

Total Solids (wt%) 0.2968 Insoluble Solids (wt%) 9.4% Temperature (°C) 32

Temperature (°C) 190 Temperature (°C) 32 Pressure (atm) 1.0

Pressure (atm) 11.4 Pressure (atm) 1.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 5 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.97

 Cellulose to Glucolig 0 Cellulose to Glucolig 0 Glucose to Zymo 0.01

 Cellulose to Cellobiose 0 Cellulose to Cellobiose 0 Glucose to Glycerol 0.002

 Cellulose to Glucose 0.0626 Cellulose to Glucose 0.9109 Glucose to Succinic Acid 0.002

 Cellulose to HMF 0 Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.005

 Xylan to Oligomer 0.0265 Glucolig to Glucose 0 Glucose to Lactic Acid 0.001

 Xylan to Xylose 0.8249 Cellobiose to Glucose 0 Xylose to Ethanol 0

 Xylan to Furfural 0 Xylan to Oligomer 0 Xylose to Zymo 0

 Xylan to Tar 0 Xylan to Xylose 0.5713 Xylose to Glycerol 0

 Mannan to Oligomer 0.0265 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0

 Mannan to Mannose 0.8249 Xylan to Tar 0 Xylose to Succinic Acid 0

 Mannan to HMF 0 Arabinan to Oligomer 0.5713 Xylose to Acetic Acid 0

 Galactan to Oligomer 0.0265 Arabinan to Arabinose 0 Xylose to Lactic Acid 0

 Galactan to Galactose 0.8249 Galactan to Oligomer 0 Arabinose to Ethanol 0

 Galactan to HMF 0 Galactan to Galactose 0.5713 Arabinose to Zymo 0

 Arabinan to Oligomer 0.0265 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Arabinose 0.8249 Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Arabinan to Furfural 0 Mannan to Mannose 0.5713 Arabinose to Acetic Acid 0

 Arabinan to Tar 0 Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Acetate to Oligomer 0 Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 Acetate to Acetic Acid 1 Soluble Sugars From PT 10,676 Galactose to Zymo 0

 Furfural to Tar 1 Other Soluble Solids From PT 20,699 Galactose to Glycerol 0

 HMF to Tar 1 Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.1 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0

S/L Separation Soluble Sugars in Produced Cellulase 0 Galactose to Lactic Acid 0

Water/Hydrolyzate Ratio (kg/kg) 0.58 Other Sol Solids in Produced Cellulase 0 Mannose to Ethanol 0

Water to S/L Separator (kg/hr) 137,034 Mannose to Zymo 0
73

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

Dilution Water (kg/hr) 67,796 Soluble Sugars From Sacc 39,360 Mannose to Glycerol 0

Conditioning Other Soluble Solids From Sacc 20,699 Mannose to Succinic Acid 0

Ca(OH)2 to Gypsum #N/A Soluble Sugars From Seed Train 0 Mannose to Acetic Acid 0

Sugar Conversion to TAR: Other Soluble Solids From Seed Train 0 Mannose to Lactic Acid 0

 Xylose 0.018 Soluble Sugars From DAP 0

 Arabinose 0.018 Other Soluble Solids From DAP 0 Contamination Loss 7.0%

 Glucose 0.006 Soluble Sugars From CSL 0

 Galactose 0.006 Other Soluble Solids From CSL 2,371 Ethanol Out of Fermenters (wt%) 4.3%

 Mannose 0.006

 Cellobiose 0.006
74

Table F-7. Operating Summary for Pervaporation Two-Stage Dilute Acid Pretreatment Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0190 Enzyme Loading (mg/g cell) 33.5 Total Solids (wt%) 20.1%

Acid Loading (g acid/g dry biomass) 0.0443 Total Solids (wt%) 20.0% Insoluble Solids (wt%) 4.2%

Total Solids (wt%) 0.2960 Insoluble Solids (wt%) 10.3% Temperature (°C) 32

Temperature (°C) 190 Temperature (°C) 32 Pressure (atm) 1.0

Pressure (atm) 11.4 Pressure (atm) 1.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 5 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.95

 Cellulose to Glucolig 0 Cellulose to Glucolig 0 Glucose to Zymo 0.02

 Cellulose to Cellobiose 0 Cellulose to Cellobiose 0 Glucose to Glycerol 0.004

 Cellulose to Glucose 0.0626 Cellulose to Glucose 0.9109 Glucose to Succinic Acid 0.006

 Cellulose to HMF 0 Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.015

 Xylan to Oligomer 0.0265 Glucolig to Glucose 0 Glucose to Lactic Acid 0.002

 Xylan to Xylose 0.8249 Cellobiose to Glucose 0 Xylose to Ethanol 0.756

 Xylan to Furfural 0 Xylan to Oligomer 0 Xylose to Zymo 0.019

 Xylan to Tar 0 Xylan to Xylose 0.5713 Xylose to Glycerol 0.003

 Mannan to Oligomer 0.0265 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0.046

 Mannan to Mannose 0.8249 Xylan to Tar 0 Xylose to Succinic Acid 0.009

 Mannan to HMF 0 Arabinan to Oligomer 0.5713 Xylose to Acetic Acid 0.014

 Galactan to Oligomer 0.0265 Arabinan to Arabinose 0 Xylose to Lactic Acid 0.002

 Galactan to Galactose 0.8249 Galactan to Oligomer 0 Arabinose to Ethanol 0

 Galactan to HMF 0 Galactan to Galactose 0.5713 Arabinose to Zymo 0

 Arabinan to Oligomer 0.0265 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Arabinose 0.8249 Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Arabinan to Furfural 0 Mannan to Mannose 0.5713 Arabinose to Acetic Acid 0

 Arabinan to Tar 0 Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Acetate to Oligomer 0 Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 Acetate to Acetic Acid 1 Soluble Sugars From PT 22,739 Galactose to Zymo 0

 Furfural to Tar 1 Other Soluble Solids From PT 18,576 Galactose to Glycerol 0

 HMF to Tar 1 Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.1 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0

S/L Separation Soluble Sugars in Produced Cellulase 0 Galactose to Lactic Acid 0

Water/Hydrolyzate Ratio (kg/kg) 0.58 Other Sol Solids in Produced Cellulase 0 Mannose to Ethanol 0

Water to S/L Separator (kg/hr) 137,371 Mannose to Zymo 0
75

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

Dilution Water (kg/hr) 45,921 Soluble Sugars From Sacc 46,152 Mannose to Glycerol 0

Conditioning Other Soluble Solids From Sacc 16,718 Mannose to Succinic Acid 0

Ca(OH)2 to Gypsum 0.174 Soluble Sugars From Seed Train 839 Mannose to Acetic Acid 0

Sugar Conversion to TAR: Other Soluble Solids From Seed Train 1,960 Mannose to Lactic Acid 0

 Xylose 0.13 Soluble Sugars From DAP 0

 Arabinose 0.2 Other Soluble Solids From DAP 0 Contamination Loss 7.0%

 Glucose 0.12 Soluble Sugars From CSL 0

 Galactose 0.28 Other Soluble Solids From CSL 2,138 Ethanol Out of Fermenters (wt%) 4.8%

 Mannose 0

 Cellobiose 0.36
76

Table F-8. Operating Summary for On-site Enzyme Production Processes

Pretreatment Saccharification Fermentation

Acid Conc (wt%) 0.0190 Enzyme Loading (mg/g cell) 0.0 Total Solids (wt%) 20.1%

Acid Loading (g acid/g dry biomass) 0.0442 Total Solids (wt%) 20.0% Insoluble Solids (wt%) 4.7%

Total Solids (wt%) 0.2959 Insoluble Solids (wt%) 10.2% Temperature (°C) 32

Temperature (°C) 190 Temperature (°C) 32 Pressure (atm) 1.0

Pressure (atm) 11.4 Pressure (atm) 1.0 Residence Time (days) 2

Residence Time (min.) Residence Time (days) 5 Conversions:

Conversions: Conversions: Glucose to Ethanol 0.95

Cellulose to Glucolig 0 Cellulose to Glucolig 0 Glucose to Zymo 0.02

 Cellulose to Cellobiose 0 Cellulose to Cellobiose 0 Glucose to Glycerol 0.004

 Cellulose to Glucose 0.0626 Cellulose to Glucose 0.9109 Glucose to Succinic Acid 0.006

 Cellulose to HMF 0 Glucolig to Cellobiose 0 Glucose to Acetic Acid 0.015

 Xylan to Oligomer 0.0265 Glucolig to Glucose 0 Glucose to Lactic Acid 0.002

 Xylan to Xylose 0.8249 Cellobiose to Glucose 0 Xylose to Ethanol 0.756

 Xylan to Furfural 0 Xylan to Oligomer 0 Xylose to Zymo 0.019

 Xylan to Tar 0.085 Xylan to Xylose 0.5713 Xylose to Glycerol 0.003

 Mannan to Oligomer 0.0265 Xylose Oligomer to Xylose 0 Xylose to Xylitol 0.046

 Mannan to Mannose 0.8249 Xylan to Tar 0 Xylose to Succinic Acid 0.009

 Mannan to HMF 0 Arabinan to Oligomer 0.5713 Xylose to Acetic Acid 0.014

 Galactan to Oligomer 0.0265 Arabinan to Arabinose 0 Xylose to Lactic Acid 0.002

 Galactan to Galactose 0.8249 Galactan to Oligomer 0 Arabinose to Ethanol 0

 Galactan to HMF 0 Galactan to Galactose 0.5713 Arabinose to Zymo 0

 Arabinan to Oligomer 0.0265 Galactose Oligomer to Galactose 0 Arabinose to Glycerol 0

 Arabinan to Arabinose 0.8249 Mannan to Oligomer 0 Arabinose to Succinic Acid 0

 Arabinan to Furfural 0 Mannan to Mannose 0.5713 Arabinose to Acetic Acid 0

 Arabinan to Tar 0.085 Mannose Oligomer to Mannose 0 Arabinose to Lactic Acid 0

 Acetate to Oligomer 0 Sugar & Solids Flow Rates (kg/hr) Galactose to Ethanol 0

 Acetate to Acetic Acid 1 Soluble Sugars From PT 20,317 Galactose to Zymo 0

 Furfural to Tar 1 Other Soluble Solids From PT 16,259 Galactose to Glycerol 0

 HMF to Tar 1 Soluble Sugars in Purchased Cellulase 0 Galactose to Succinic Acid 0

 Lignin to Soluble Lignin 0.1 Other Sol Solids in Purchased Cellulase 0 Galactose to Acetic Acid 0

S/L Separation Soluble Sugars in Produced Cellulase 545 Galactose to Lactic Acid 0

Water/Hydrolyzate Ratio (kg/kg) 0.58 Other Sol Solids in Produced Cellulase 2,781 Mannose to Ethanol 0

Water to S/L Separator (kg/hr) 137,416 Mannose to Zymo 0
77

Pretreatment, cont. Saccharification, cont. Fermentation, cont.

Dilution Water (kg/hr) 19,186 Soluble Sugars From Sacc 41,246 Mannose to Glycerol 0

Conditioning Other Soluble Solids From Sacc 17,136 Mannose to Succinic Acid 0

Ca(OH)2 to Gypsum 0.175 Soluble Sugars From Seed Train 783 Mannose to Acetic Acid 0

Sugar Conversion to TAR: Other Soluble Solids From Seed Train 2,003 Mannose to Lactic Acid 0

 Xylose 0.13 Soluble Sugars From DAP 0

 Arabinose 0.2 Other Soluble Solids From DAP 0 Contamination Loss 7.0%

 Glucose 0.12 Soluble Sugars From CSL 0

 Galactose 0.28 Other Soluble Solids From CSL 2,056 Ethanol Out of Fermenters (wt%) 4.5%

 Mannose 0

 Cellobiose 0.36
78

Appendix G - General Process Description

Area 100

The feedstock handling area receives the corn stover in bales. After being unwrapped, stover is 

washed to remove dirt before being passed through a magnetic separator to remove tramp metal 

and then conveyed to the shredders for size reduction.

Area 200

In the pretreatment area, biomass undergoes a physical and/or chemical treatment, which allows 

for improved exposure of the cellulose during enzymatic hydrolysis. The details of the four 

pretreatment processes are discussed under the Pretreatment Variation section of process 

variations in the results section.

Area 300

Enzymatic saccharification followed by fermentation of sugars occurs in this area. The 

pretreatment hydrolyzate from Area 200 is pumped into 18 parallel saccharification vessels (each 

has a capacity of 1 MM gal) where enzyme is added (enzyme loading of 31.3 mg protein/g 

cellulose) either from purchased stock preparation or from the on-site enzyme production section 

(Area 400). The residence time for saccharification is 5 days. Note that for two-stage dilute acid 

treatment, acid is used in place of enzyme to hydrolyze the cellulose and xylan. 

A small fraction of the hydrolyzate leaving the saccharification vessel is sent to one of two 

sequenced batch fermentation seed vessel trains to be used as a carbon source for the growth of 

the fermentative organism, Zymomonas mobilis. Nutrients such as CSL and diammonium 

phosphate are also added into the bio-reactors. 

The bulk of the hydrolyzate from the saccharification vessels is pumped to one of eight parallel 

sequenced batch fermentation reactors, where Zymomonas mobilis is introduced from the seed 

reactors. A 2 day residence time is used for fermentation. Vent gas from the fermentors is 

scrubbed to collect escaped ethanol in the scrubber and sent to the beer column along with 

fermented beer from the fermentors. 

Area 500

Ethanol is separated and recovered from water and residual solids in this section. Fermented beer 

is distilled in two distillation columns: beer column and rectification column. The beer column 

removes dissolved CO2 and vapor generating a concentrated ethanol stream. The rectification 

column purifies the ethanol to about 95 wt% concentration. Nearly all of the remaining water is 

removed in a molecular sieve adsorption column. The stillage from the beer column is pumped to 

the 1st-effect evaporator where water content is reduced, followed by dewatering in a 

Pneumapress filter and screw press to separate most of the insoluble solids, comprising mainly

lignin, from the aqueous stream. The liquid fraction from the screw press is concentrated in the 

2nd- and the 3rd-effect evaporators to high concentration syrup of soluble solids. The evaporator 

syrup and the dewatered solid fraction from the screw press are used as boiler fuel in the 

fluidized bed combustor. 
79

Boiler Blowdown

Evaporator Condensate

(435 MT/d)

(155 MT/d)

Equalizing 

Basin

T = 35 oC

P = 1 atm

To Scrubber

(7 MT/d)

Nutrient 

Feed 

System

(477 MT/d)

Pretreatment Flash

(1073 MT/d)

T = 35 o

C

P = 1 atm Cooling Tower Blowdown

(379 MT/d)

S/L Sep Vent Condensate

(12 MT/d)

Aerobic 

Digester

Air Vapor to 

Atmosphere

(485 MT/d)

Aerobic Digester

Nutrient

(0.7 MT/d) Wastewater

Wastewater

(2048 MT/d)

Clarifier

T = 21.1 oC

P = 1 atm

To Process 

Water Inlet

(2037 MT/d)

Sludge To 

Combustor

(2.7 MT/d)

S/L 

Separation

Water Return

(14 MT/d)

(16 MT/d)

Solid Suspension

To Clarifier

(2053 MT/d)

Area 600

Wastewater is treated in anaerobic and aerobic digesters. In the anaerobic digester, a small 

amount of urea, phosphoric acid, and micronutrients are added as nutrients for the anaerobic 

organisms. Biogas from the anaerobic digester is used as boiler fuel. The wastewater is further 

treated aerobically. It is then held in a clarifying tank where the settled solids are separated from 

the water. The solids are dewatered in a belt filter press, with a polymer being added to aid in 

dewatering, followed by a screw press. The resulting sludge is used as boiler fuel. The water 

from the clarifying tank is recycled as process water. 

 

Figure G-1. Wastewater treatment section (Area 600)
80

Lignin (1137 MT/d)

To Deareator

Steam to Evaporator

Evaporator Syrup

(1882 MT/d)

Digester Solids (3 MT/d)

Biogas (7 MT/d)

Boiler Feed Water (5181 MT/d)

Preheated Air (7794 MT/d)

Blowdown to WWT 

(155 MT/d)

Steam to Area 200, 

300, 500, 800 & 900

Flue Gas

Generator

To Plant

To Grid

Super heated Steam

Boiler

Turbine

Ash (279 MT/d)

To Stack

(10544 MT/d) Baghouse

To Pretreatment Reactor

(906 MT/d)

(2361 MT/d)

(129 MT/d)

(5026 MT/d)

(1630 MT/d)

Area 800

Cogeneration of steam for process heat and electricity occurs in Area 800. Evaporator syrup, 

insoluble solids from the Pneumapress, wastewater treatment sludge, and biogas are used as fuel 

in the fluidized bed combustor. Superheated steam is generated in the boiler and exits the multistage turbine at three different conditions needed in the process. 

Figure G-2. Steam and power generation section (Area 800)
81

Appendix H - Cost By Area Of Process Scenarios

Table H-1. Costs by Area of the Dilute Acid Pretreatment (High Solids) Scenario

Cost Areas / Factor

Installed Cost Purchased 

Equipment Cost

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 6.4 6.0 5.3

Pretreatment (Area 200) 38.0 22.4 21.0 18.5

Saccharification & Fermentation (Area 300) 21.8 12.9 17.3 15.3

Distillation and Solids Recovery (Area 500) 25.7 15.1 16.9 14.9

Wastewater Treatment (Area 600) 5.8 3.4 4.5 4.0

Storage (Area 700) 3.1 1.8 1.9 1.7

Boiler/Turbogenerator (Area 800) 57.6 24.4 41.4 36.5

Utilities (Area 900) 6.6 3.9 4.3 3.8

Purchased Equipment Cost 113.3 100

Total Installed Equipment Cost 169.4 100

Fixed Capital Investment (FCI) 337.8

Working Capital (WC) 50.7

Total Capital Investment (TCI) 388.5

Lang Factor 3.43

Table H-2. Costs by Area of the 2-Stage Dilute Acid Pretreatment Scenario 

Cost Areas / Factor

Installed Cost Purchased 

Equipment Cost 

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 6.3 6.0 5.4

Pretreatment (Area 200) 44.9 26.0 23.5 21.0

Saccharification & Fermentation (Area 300) 9.7 5.6 7.2 6.4

Distillation and Solids Recovery (Area 500) 26.7 15.5 18.1 16.1

Wastewater Treatment (Area 600) 4.6 2.7 3.5 3.1

Storage (Area 700) 2.8 1.6 1.7 1.5

Boiler/Turbogenerator (Area 800) 66.2 38.3 47.6 42.5

Utilities (Area 900) 6.9 4.0 4.5 4.0

Purchased Equipment Cost 112.1 100

Total Installed Equipment Cost 172.7 100

Fixed Capital Investment (FCI) 340.0

Working Capital (WC) 51.0

Total Capital Investment (TCI) 391.0

Lang Factor 3.49
82

Table H-3. Costs by Area of the Hot Water Pretreatment Scenario 

Cost Areas / Factor

Installed Cost Purchased Equipment 

Cost 

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 7.0 6.0 5.6

Pretreatment (Area 200) 6.7 4.3 3.9 3.6

Saccharification & Fermentation (Area 300) 30.2 19.3 23.4 21.7

Distillation and Solids Recovery (Area 500) 30.9 19.8 19.4 18.0

Wastewater Treatment (Area 600) 1.9 1.2 1.3 1.2

Storage (Area 700) 3.3 2.1 2.1 2.0

Boiler/Turbogenerator (Area 800) 65.8 42.1 47.3 43.9

Utilities (Area 900) 6.7 4.3 4.3 4.0

Purchased Equipment Cost 107.8 100

Total Installed Equipment Cost 156.3 100

Fixed Capital Investment (FCI) 314.0

Working Capital (WC) 47.1

Total Capital Investment (TCI) 361.1

Lang Factor 3.35

Table H-4. Costs by Area of the AFEX Pretreatment Scenario 

Cost Areas / Factor

Installed Cost Purchased Equipment 

Cost 

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 6.5 6.0 5.3

Pretreatment (Area 200) 30.8 18.4 18.4 16.1

Saccharification & Fermentation (Area 300) 23.5 14.0 18.8 16.5

Distillation and Solids Recovery (Area 500) 27.5 16.4 17.5 15.4

Wastewater Treatment (Area 600) 1.6 1.0 1.1 1.0

Storage (Area 700) 2.8 1.7 1.8 1.6

Boiler/Turbogenerator (Area 800) 62.0 37.0 44.6 39.1

Utilities (Area 900) 8.4 5.0 5.8 5.1

Purchased Equipment Cost 113.9 100

Total Installed Equipment Cost 167.4 100

Fixed Capital Investment (FCI) 335.6

Working Capital (WC) 50.4

Total Capital Investment (TCI) 386.0

Lang Factor 3.39
83

Table H-5. Costs by Area of the Separate C-5 and C-6 Fermentation Process Scenario 

Cost Areas / Factor

Installed Cost Purchased Equipment 

Cost 

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 11.4 6.8 6.2 5.5

Pretreatment (Area 200) 34.2 20.4 18.4 16.4

Saccharification & Fermentation (Area 300) 29.6 17.6 23.7 21.1

Distillation and Solids Recovery (Area 500) 27.3 16.3 17.8 15.8

Wastewater Treatment (Area 600) 3.1 1.8 2.3 2.0

Storage (Area 700) 5.1 3.0 3.4 3.0

Boiler/Turbogenerator (Area 800) 50.9 30.3 36.5 32.5

Utilities (Area 900) 6.2 3.7 4.1 3.6

Purchased Equipment Cost 112.4 100

Total Installed Equipment Cost 167.8 100

Fixed Capital Investment (FCI) 335.4

Working Capital (WC) 50.3

Total Capital Investment (TCI) 385.8

Lang Factor 3.43

Table H-6. Costs by Area of the Pervaporation Process Scenario 

Cost Areas / Factor

Installed Cost Purchased Equipment 

Cost 

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 5.2 6.0 3.9

Pretreatment (Area 200) 36.2 17.3 19.7 12.7

Saccharification & Fermentation (Area 300) 21.9 10.5 17.4 11.2

Distillation and Solids Recovery (Area 500) 70.8 33.8 62.9 40.5

Wastewater Treatment (Area 600) 3.3 1.6 2.4 1.5

Storage (Area 700) 3.3 1.6 2.1 1.4

Boiler/Turbogenerator (Area 800) 55.9 26.7 40.1 25.8

Utilities (Area 900) 7.0 3.3 4.6 3.0

Purchased Equipment Cost 155.2 100

Total Installed Equipment Cost 209.2 100

Fixed Capital Investment (FCI) 435.7

Working Capital (WC) 65.4

Total Capital Investment (TCI) 501.0

Lang Factor 3.23
84

Table H-7. Costs by Area of the On-site Enzyme Production Processes Scenario 

Cost Areas / Factor

Installed Cost Purchased Equipment 

Cost 

(MM$) (%) (MM$) (%)

Feedstock Handling (Area 100) 10.9 5.8 6.0 4.7

Pretreatment (Area 200) 36.2 19.3 19.7 15.5

Saccharification & Fermentation (Area 300) 21.8 11.6 17.3 13.6

On-site Enzyme Production (Area 400) 23.7 12.6 17.9 14.0

Distillation and Solids Recovery (Area 500) 25.3 13.5 16.8 13.2

Wastewater Treatment (Area 600) 3.4 1.9 2.6 2.0

Storage (Area 700) 2.5 1.3 1.6 1.2

Boiler/Turbogenerator (Area 800) 57.2 30.4 41.1 32.3

Utilities (Area 900) 6.8 3.6 4.4 3.5

Purchased Equipment Cost 127.4 100

Total Installed Equipment Cost 187.8 100

Fixed Capital Investment (FCI) 377.5

Working Capital (WC) 56.6

Total Capital Investment (TCI) 434.2

Lang Factor 3.41
85

Appendix I - Down Selection Matrix

Table I-1. Process Down-Selection Matrix

Capital Expenditure (x106$)

Operating Costs (x106$)

Fuel Cost ($/gal)Internal Rate of Return (%)RiskPlant Efficiency (%)Carbon EfficiencyProcess Yield (fuel) [gal Ethanol or Butanol/ ton feed stock]Capacity FactorOptimal Plant Size (MTon/day)Complexity of ProcessTechnology Development

Water Use/Discharge (Mton/day)Greenhouse Gas EmissionsParticulates/Solid WastesAir EmissionsWaste TreatmentToxicsEnergy Content (Gasoline Equivalent)Engine CompatibilityUse EmissionsInfrastructure CompatibilityFuel ToxicityAbility to Meet Fuel StandardsBibliographya

Notes

Process Blocks

Ethanol Concentrated 

Acid/Fermentation

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

3.24 0.16 1.07 10 48 89.7 93 2000

Pilot

5440 0.658 1

Process Blocks: (1) Pretreatment 

(dilute H2SO4); (2) Continuous 

Saccharification; and (3) 

Cofermentation (hydrolysis & 

fermentation units are in series); (4) 

Separation/Purification (distillation and 

molecular sieve); (5) Water treatment 

(treated water recycle to process) 

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

2.63 10 91 2000

Commercial/Pilot

0.628 2

Scenario-I 

(Short Term: 5 

years, following 

pretreatment 

and SSF)

Process Blocks: (1) Pretreatment 

(dilute acid); (2) Hydrolysis 

(Enzymatic); (3) Fermentation; (4) 

Separation/Purification (distillation & 

desiccants)

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

4.53 1.71 12 48 76.2 95.9 5000

Commercial/Pilot

3

Current 

Process 

(pretreatment & 

SSF)

Process Blocks: (1) Pretreatment 

(dilute acid); (2a) SSF (simultaneous 

sccharification & fermentation); (2b) 

with separate fermentation for C5; (3) 

Separation/Purification (2-column 

distillation); (4) Water treatment 

(suspended sludge digester)

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

1.27 1.02 0.88 50 119.5 90.4 634 0.658 4

Feedstock: 

Corn

Process Blocks: (1) Pretreatment 

(dilute H2SO4); (2) Simultaneous 

Saccharification and Cofermentation

(glucose & Xylose); and (3) 

Separation/Purification (distillation and 

molecular sieve); (5) Water treatment 

(treated water recycle to process) 

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

4.83 1.68 1.45 42 79.2 90.4 957 0.658 5

Feedstock: 

Cornstover

Process Blocks: (1) Pretreatment 

(dilute H2SO4); (2) Simultaneous 

Saccharification and Cofermentation

(glucose & Xylose); and (3) 

Separation/Purification (distillation and 

molecular sieve); (5) Water treatment 

(treated water recycle to process) 

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

4.91 0.25 1.14 10 36 68 95.9 2000

Pilot

6307 0.661 6

SSCF Base 

Case; 

Feedstock: 

Yellow poplar

Process Blocks: (1) Pretreatment 

(dilute H2SO4); (2) Simultaneous 

sccharification and cofermentation (of 

glucose & xylan); (3) 

Separation/Purification; (5) Water 

treatment (treated water recycle to 

process) 
86

Capital Expenditure (x106$)

Operating Costs (x106$)

Fuel Cost ($/gal)Internal Rate of Return (%)RiskPlant Efficiency (%)Carbon EfficiencyProcess Yield (fuel) [gal Ethanol or Butanol/ ton feed stock]Capacity FactorOptimal Plant Size (MTon/day)Complexity of ProcessTechnology Development

Water Use/Discharge (Mton/day)Greenhouse Gas EmissionsParticulates/Solid WastesAir EmissionsWaste TreatmentToxicsEnergy Content (Gasoline Equivalent)Engine CompatibilityUse EmissionsInfrastructure CompatibilityFuel ToxicityAbility to Meet Fuel StandardsBibliographya

Notes

Process Blocks

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

3.85 1.16 10 40 76 95.9 2000

Pilot

6307 0.6617

SSCF Near 

Term; 

Feedstock: 

Yellow poplar

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

2.98 0.94 10 42 81 95.9 2000

Pilot

6307 0.6618

SSCF 2005; 

Feedstock: 

Yellow poplar

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

2.37 0.82 10 49 94 95.9 2000

Pilot

6307 0.6619

SSCF 2010; 

Feedstock: 

Yellow poplar

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

2.29 0.76 10 52 99 95.9 2000

Pilot

6307 0.661 10

SSCF 2015; 

Feedstock: 

Yellow poplar

Ethanol Dilute Acid 

Pretreatment/Enzymatic 

/ Fermentation

8.67 1.76 1.25 15 78 33 100 4360

Pilot

11

SHF (TVA 

process); High 

profit from co-

products: 

furfural

Process Blocks: (1) Pretreatment 

(dilute H

2SO4); (2) Hydrolysis 

(Enzymatic); (3) Fermentation 

(Hexose & Pentose in separate 

vessels); (4) Separation/Purification 

(distillation); (5) Waste treatment

Ethanol SO

2

-Steam 

Pretreatment/Enzymatic 

/ Fermentation

5.58 0.45 0.98 25 42 78.5 100 500 12

SO

2 & NaOH in 

2-step 

preteratment; 

Rate of return 

on working 

capital

Process Blocks: (1) Pretreatment 

(SO

2, Steam); (2) Hydrolysis 

(Enzymatic); (3a) Fermentation, and 

(3b) fermentation of C5 sugars in 

separate vessels; (4) Seperation; (5) 

water treatment

Ethanol SO

2

-Steam 

Pretreatment/Enzymatic 

/ Fermentation

13.61 0.755 36 69.6 91.2 586.8

lab

552 0.624 13

SHF process; 

SO

2 is used for 

pretreatment; 

Total 

Investment is 

1184 MSKR

Process Blocks: (1) Pretreatment 

(Stake Tech: SO

2 & Steam); (2) 

Hydrolysis (Enzymatic); (3) 

Fermentation; (4) 

Separation/Purification (distillation & 

drying)

Ethanol SO

2

-Steam 

Pretreatment/Enzymatic 

/ Fermentation

9.18 0.835 41 79.2 91.2 586.8

lab

586 0.618 14

SSF process; 

SO

2 is used for 

pretreatment; 

Total 

Investment is 

910 MSKR

Process Blocks: (1) Pretreatment 

(Stake Tech: SO2 & Steam); (2) SSF 

(Simultaneous Saccharification & 

Fermentation); (3) 

Separation/Purification (distillation & 

drying)

Ethanol SO

2

-Steam 

Pretreatment/Enzymatic 

/ Fermentation

8.63 2.246 40 76.8 91.2 600

lab

0.621 15

SSF process; 

SO

2 is used for 

pretreatment; 

One-Step 

Pretreatment; 

Total 

Investment is 

847 MSKR

Process Blocks: (1) Pretreatment 

(Stake Tech: SO2 & Steam); (2) SSF 

(Simultaneous Saccharification & 

Fermentation; feed batch); (3) 

Separation/Purification (2-separate 

stripper & rectifier operations)
87

Capital Expenditure (x106$)

Operating Costs (x106$)

Fuel Cost ($/gal)Internal Rate of Return (%)RiskPlant Efficiency (%)Carbon EfficiencyProcess Yield (fuel) [gal Ethanol or Butanol/ ton feed stock]Capacity FactorOptimal Plant Size (MTon/day)Complexity of ProcessTechnology Development

Water Use/Discharge (Mton/day)Greenhouse Gas EmissionsParticulates/Solid WastesAir EmissionsWaste TreatmentToxicsEnergy Content (Gasoline Equivalent)Engine CompatibilityUse EmissionsInfrastructure CompatibilityFuel ToxicityAbility to Meet Fuel StandardsBibliographya

Notes

Process Blocks

Ethanol SO

2

-Steam 

Double 

Pretreatment/Enzymatic 

/ Fermentation

8.72 2.246 42 80.1 91.2 600

lab

0.621 16

SSF process; 

SO

2 is used for 

pretreatment; 

Two-Step 

Pretreatment; 

Total 

Investment is 

847 MSKR

Process Blocks: (1) Pretreatment (2-

step SO

2

-steam pretreatment: double 

pretreatment); (2) SSF (Simultaneous 

Saccharification & Fermentation; feed 

batch); (3) Separation/Purification (2-

separate stripper & rectifier 

operations)

Ethanol AFEX 

Pretreatment/Enzymatic 

/ Fermentation

1.95 0.73 12 61 105.3 95.9 5000

Lab

17

Advanced 

EtOH Rankine 

process

Process Blocks: (1) Pretreatment 

(Ammonia fiber expansion, AFEX, 

pretreatment - no detoxification); (2) 

SSF (simultaneous sccharification & 

fermentation); (3) 

Separation/Purification (1-column 

distillation with heat integration); (4) 

Water treatment (attached film 

digester)

Ethanol AFEX 

Pretreatment/Enzymatic 

/ Fermentation

2.89 0.77 12 61 105.3 95.9 5000

conceptual

18

Advanced 

EtOH-GTCC 

Process

Process Blocks: (1) Pretreatment 

(Ammonia fiber expansion, AFEX, 

pretreatment- no detoxification); (2) 

CBP (consolidated bio-processing); (3) 

Separation/Purification (1-column 

distillation with heat integration); (4) 

Water treatment (attached film 

digester)

E thanol Liquid 

Hotwater 

Pretreatment/Enzymatic 

/ Fermentation

0.99 10 91 2000

conceptual

0.628 19

Scenario-III 

(Long Term: 

>20 years, 

following 

Consolidated 

BioProcessing, 

CBP)

Process Blocks: (1) Pretreatment 

(Liquid Hot Water); (2) Consolidated 

Bio-Processing; (3) 

Separation/Purification (distillation & 

desiccants)

Ethanol Steam 

Pretreatment/Enzymatic 

/ Fermentation

681-

907 1.48 10 91 2000

lab/pilot

0.628 20

Scenario-II 

(Mid Term: 5-

10 years, 

following 

pretreatment 

and SSCF)

Process Blocks: (1) Pretreatment 

(steam explosion); (2) SSF & 

SSCF(Simultaneous saccharification &

fermentation, and simultaneous 

saccharification and cofermentation); 

(3) Separation/Purification (distillation 

& desiccants)

Ethanol Enzymatic / 

Fermentation

9.06 1.33 1.1 15 47 79.4 100 15

Lab

0.658 21

Scenario-I 

(minerals sent 

to land fill; No 

ethanol from 

xylan)

Process Blocks: (1) SSF 

(Simultaneous saccharification and 

fermentation); (2) 

Separation/Purification (distillation and 

molecular sieve); (3) Water treatment 

Ethanol Enzymatic / 

Fermentation

7.70 0.85 1.1 15 55 93.5 100 15

no-lab demo

0.658 22

Scenario-II 

(90% of 

minerals are 

recovered and 

sent to paper 

mill)

Process Blocks: (1) SSF 

(Simultaneous saccharification and 

fermentation); (2) 

Separation/Purification (distillation and 

molecular sieve); (3) Water treatment 

Butanol Concentrated 

Acid / Fermentation

Butanol Dilute Acid 

Pretreatment / 

Fermentation

44 28.4

lab

0.957 23

Feedstock: 

corn fiber; 

dilute H

2SO4

hydrolysis

Process Blocks: (1) Hydrolysis (dilute H2SO4); (2) Fermentation; (3) 

Separation/Purification (distillation)
88

Capital Expenditure (x106$)

Operating Costs (x106$)

Fuel Cost ($/gal)Internal Rate of Return (%)RiskPlant Efficiency (%)Carbon EfficiencyProcess Yield (fuel) [gal Ethanol or Butanol/ ton feed stock]Capacity FactorOptimal Plant Size (MTon/day)Complexity of ProcessTechnology Development

Water Use/Discharge (Mton/day)Greenhouse Gas EmissionsParticulates/Solid WastesAir EmissionsWaste TreatmentToxicsEnergy Content (Gasoline Equivalent)Engine CompatibilityUse EmissionsInfrastructure CompatibilityFuel ToxicityAbility to Meet Fuel StandardsBibliographya

Notes

Process Blocks

Butanol AFEX 

Pretreatment / 

Enzymatic / 

Fermentation

Butanol Hotwater 

Pretreatment / 

Enzymatic / 

Fermentation

65 57.6

Lab

0.957 24

Feedstock: 

Saccharified 

Liquefied 

Cornstarch 

(SLCS) with 

moisture 

content of 

approximately 

60% 

Process Blocks: (1) Hydrolysis 

(enzymatic hydrolysis of LCS to 

SLCS); (2) Fermentation; (3) 

Separation/Purification

Butanol Hotwater 

Pretreatment / 

Fermentation

60 54

Lab

0.957 25

Feedstock: 

Liquefied 

Cornstarch 

(LCS) with 

moisture 

content of 

approximately 

60% 

Process Blocks: (1) Simultaneous 

saccharification and Fermentation; (2) 

Separation/Purification

Butanol Enzyme / 

Fermentation

0.81 0.07 0.74

10-

30 66 89.84 95.9 1469

lab

0.949 26

Feedstock: 

corn (14% 

moisture); 

Hydrolysis + 

Immobilized 

Cell 

Continuous 

Fermentation & 

pervaporative 

recovery, 

ICCFPR 

process

Process Blocks: (1) Hydrolysis 

(enzymatic); (2) Fermentation 

(contineous); (3) 

Separation/Purification (pervaporation 

& distillation)

Butanol Fermentation

1.31 0.15 0.81 15 58 5.52 100 2268

lab

0.957 27

Feedstock: 

Whey permeat 

(4.5% lactose); 

Fibrous Bed 

Reactor, FBR; 

Process: 2 step 

fermentation

Process Blocks: (1) 2-step 

Fermentation (i. fermentation of 

lactose to butyric acid, ii. Fermentation 

of butyric acis to butanil); (3) 

Separation/Purification (adsorption, 

desorption, distillation); 

Butanol Fermentation

57 114

lab

0.957

Feedstock: 

Glucose; 2-

step 

fermentation, 

FBR

Process Blocks: (1) hydrolysis of 

lignocellulosics/starch to fermentable 

sugars; (2) 2-step Fermentation of 

glucose (i. fermentation of lactose to 

butyric acid, ii. Fermentation of butyric 

acis to butanil); (3) 

Separation/Purification (adsorption, desorption, distillation); 

Butanol Fermentation

2.37 0.22 1.68 20 65 89.79 95.9 1469

Pilot

0.947 28

Feedstock: 

corn (14% 

moisture)

Process Blocks: (1) Simultaneous 

saccharification and Fermentation; (2) 

Separation/Purification (distillation); (3) 

Wastewater treatment (recycling to 

process) 

Butanol Fermentation

1.00 0.09 0.95

10-

30 66 89.84 95.9 1469

Pilot

0.949 29

Feedstock: 

corn (14% 

moisture); 

Batch 

Fermentation & 

pervaporative 

recovery, 

BFPR process

Process Blocks: (1) Simultaneous 

saccharification and Fermentation; (2) 

Separation/Purification (pervaporation 

& distillation)
89

Capital Expenditure (x106$)

Operating Costs (x106$)

Fuel Cost ($/gal)Internal Rate of Return (%)RiskPlant Efficiency (%)Carbon EfficiencyProcess Yield (fuel) [gal Ethanol or Butanol/ ton feed stock]Capacity FactorOptimal Plant Size (MTon/day)Complexity of ProcessTechnology Development

Water Use/Discharge (Mton/day)Greenhouse Gas EmissionsParticulates/Solid WastesAir EmissionsWaste TreatmentToxicsEnergy Content (Gasoline Equivalent)Engine CompatibilityUse EmissionsInfrastructure CompatibilityFuel ToxicityAbility to Meet Fuel StandardsBibliographya

Notes

Process Blocks

Butanol Fermentation

0.93 0.08 0.81

10-

30 66 89.84 95.9 1469

Pilot

0.949 30

Feedstock: 

corn (14% 

moisture); Fedbatch 

Fermentation & 

pervaporative 

recovery, 

FBFPR 

process

Process Blocks: (1) Simultaneous 

saccharification and Fermentation; (2) 

Separation/Purification (pervaporation 

& distillation)

Zeachem Process

a Bibliography references:

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90

Appendix J - Analysis: Cost Growth Variables And Results

Table J-1. Plant Performance and Cost Growth Variables for Dilute Acid Pretreatment (High 

Solids Loading) Processes

Plant Performance (Equation 1) Cost Growth (Equation 2)

Variables

Values

Variables

Values

Optimistic

Most 

Probable

PessimisticOptimisticMost 

Probable

Pessimistic

NEWSTEPSa 6 6 6 PCTNEW 61.4 61.4 61.4

BALEQS 50 40 30 IMPURITIES 0 3 5

WASTE 1 2 3 COMPLEXITYb 6 6 6

SOLIDS 1 1 1 INCLUSIVENESS 33 0 0

PROJECT 

DEFINITION 6 6 7

Plant Performance 

(%) 22.1 14.7 7.3 Cost Growth (%) 52.6 42.6 32.0

a New steps/units: Feedstock handling, Pretreatment, Saccharification, Cofermentation, Beer Column, and 

Combustor. b Continuously linked steps: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, 

Distillation, Steam/Power Generation.

Table J-2. Plant Performance and Cost Growth Variables for Hot water Pretreatment Processes

Plant Performance (Equation 1) Cost Growth (Equation 2)

Variables

Values

Variables

Values

Optimistic

Most 

Probable

PessimisticOptimisticMost 

Probable

Pessimistic

NEWSTEPSa 6 6 6 PCTNEW 53.6 53.6 53.6

BALEQS 50 40 30 IMPURITIES 0 3 5

WASTE 1 2 3 COMPLEXITYb 6 6 6

SOLIDS 1 1 1 INCLUSIVENESS 33 0 0

PROJECT 

DEFINITION 6 6 7

Plant Performance 

(%) 22.1 14.7 7.3 Cost Growth (%) 55.0 44.9 34.3

a New steps/units: Feedstock handling, Pretreatment, Saccharification, Cofermentation, Beer Column, and 

Combustor. b Continuously linked steps: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, 

Distillation, Steam/Power Generation.
91

Table J-3. Plant Performance and Cost Growth Variables for Two-Stage Dilute Acid 

Pretreatment Processes

Plant Performance (Equation 1) Cost Growth (Equation 2)

Variables

Values

Variables

Values

Optimistic

Most 

Probable

PessimisticOptimisticMost 

Probable

Pessimistic

NEWSTEPSa 6 6 6 PCTNEW 60.8 60.8 60.8

BALEQS 50 40 30 IMPURITIES 0 3 5

WASTE 1 2 3 COMPLEXITYb 6 6 6

SOLIDS 1 1 1 INCLUSIVENESS 33 0 0

PROJECT 

DEFINITION 6 6 7

Plant Performance 

(%) 22.1 14.7 7.3 Cost Growth (%) 50.8 42.8 32.2

a New steps/units: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, Beer Column, and 

Combustor. b Continuously linked steps: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, 

Distillation, Steam/Power Generation.

Table J-4. Plant Performance and Cost Growth Variables for AFEX Pretreatment Processes

Plant Performance (Equation 1) Cost Growth (Equation 2)

Variables

Values

Variables

Values

Optimistic

Most 

Probable

PessimisticOptimisticMost 

Probable

Pessimistic

NEWSTEPSa 7 7 7 PCTNEW 60.38 60.38 60.38

BALEQS 50 40 30 IMPURITIES 0 3 5

WASTE 1 2 2 COMPLEXITYb 7 7 7

SOLIDS 1 1 1 INCLUSIVENESS 33 0 0

PROJECT 

DEFINITION 6 6 7

Plant Performance 

(%) 12.4 5.0 1.7 Cost Growth (%) 51.8 41.8 31.2

a New steps/units: Feedstock Handling, Pretreatment, Saccharification, Cofermentation, Ammonia Separation, 

Beer Column, and Combustor. b Continuously linked steps: Feedstock Handling, Pretreatment, Saccharification, 

Cofermentation, Ammonia recovery, Distillation, Steam/Power Generation.
92

Table J-5. Plant Performance and Cost Growth Variables for On-Site Enzyme Production 

Processes

Plant Performance (Equation 1) Cost Growth (Equation 2)

Variables

Values

Variables

Values

Optimistic

Most 

Probable

PessimisticOptimisticMost 

Probable

Pessimistic

NEWSTEPSa 7 7 7 PCTNEW 66.9 66.9 66.9

BALEQS 50 40 30 IMPURITIES 0 3 5

WASTE 1 2 3 COMPLEXITYb 7 7 7

SOLIDS 1 1 1 INCLUSIVENESS 33 0 0

PROJECT 

DEFINITION 6 6 7

Plant Performance

(%) 12.4 5.0 0 Cost Growth (%) 49.9 39.8 29.2

a New steps/units: Feedstock handling, Pretreatment, Saccharification, Cofermentation, Beer Column, 

Combustor, and Enzyme Production. b Continuously linked steps: Feedstock Handling, Pretreatment, 

Saccharification, Cofermentation, Distillation, Steam/Power Generation, and Enzyme Production.

Table J-6. PV from Cost Growth Analysis for Various Process Scenarios 

Cost Item Cost Growth (Pioneer Plant)a

Most Probable Optimistic Pessimistic

Dilute Acid Pretreatment Processes (Pilot)

PV ($/Gal) 6.11 5.27 7.59

Total Capital Investment (MM$) 913.9 750.1 1,198.8

Lang Factor 8.06 6.62 10.58

Hot Water Pretreatment Processes

PV ($/Gal) 7.32 6.41 8.87

Total Capital Investment (MM$) 808.3 619.0 1,042.5

Lang Factor 7.50 6.21 9.67

2-Stage Dilute Acid Pretreatment Processes

PV ($/Gal) 8.36 7.10 10.56

Total Capital Investment (MM$) 916.2 752.6 1,200.2

Lang Factor 8.17 6.71 10.70

AFEX Pretreatment Processes

PV ($/Gal) 6.84 5.85 8.41

Total Capital Investment (MM$) 925.2 756.5 1,221.7

Lang Factor 8.12 6.64 10.72

On-site Enzyme Production Processes

PV ($/Gal) 7.21 6.04 9.21

Total Capital Investment (MM$) 1,088.1 881.4 1460.9

Lang Factor 8.54 6.92 11.47

a 30% contingency is used for pioneer plant cost analysis. 
F1147-E(10/2008)

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June 2010

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Technical report

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4. TITLE AND SUBTITLE

Techno-Economic Analysis of Biochemical Scenarios for Production 

of Cellulosic Ethanol

5a. CONTRACT NUMBER

DE-AC36-08-GO28308

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6. AUTHOR(S)

F. Kabir Kazi, J. Fortman, R. Anex, G. Kothandaraman, D. Hsu, 

A, Aden, and A. Dutta

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NREL/TP-6A2-46588

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BB07.7510

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National Renewable Energy Laboratory

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Golden, CO 80401-3393

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NREL/TP-6A2-46588

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14. ABSTRACT (Maximum 200 Words)

A techno-economic analysis on the production of cellulosic ethanol by fermentation was conducted to understand the 

viability of liquid biofuel production processes within the next 5-8 years. Initially, 35 technologies were reviewed, then 

a two-step down selection was performed to choose scenarios to be evaluated in a more detailed economic analysis. 

The lignocellulosic ethanol process was selected because it is well studied and portions of the process have been 

tested at pilot scales. Seven process variations were selected and examined in detail. Process designs were 

constrained to public data published in 2007 or earlier, without projecting for future process improvements. Economic 

analysis was performed for an 'nth plant' (mature technology) to obtain total investment and product value (PV). 

Sensitivity analysis was performed on PV to assess the impact of variations in process and economic parameters. 

Results show that the modeled dilute acid pretreatment process without any downstream process variation had the 

lowest PV of $3.40/gal of ethanol ($5.15/gallon of gasoline equivalent) in 2007 dollars. Sensitivity analysis shows that 

PV is most sensitive to feedstock and enzyme costs.

15. SUBJECT TERMS

cellulosic ethanol; biofuels; product value; fermentation; biochemical; feedstock; enzyme; process design; economic 

analysis; dilute acid pretreatment

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