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

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

Contract No. DE-AC36-08GO28308

Techno-Economic Analysis of 

Biofuels Production Based on 

Gasification

Ryan M. Swanson, Justinus A. Satrio, and 

Robert C. Brown

Iowa State University

Alexandru Platon

ConocoPhillips Company

David D. Hsu

National Renewable Energy Laboratory

Technical Report

NREL/TP-6A20-46587

November 2010
NREL is a national laboratory of the U.S. Department of Energy, Office of Energy 

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

National Renewable Energy Laboratory

1617 Cole Boulevard

Golden, Colorado 80401

303-275-3000 • www.nrel.gov

Contract No. DE-AC36-08GO28308

Techno-Economic Analysis of 

Biofuels Production Based on 

Gasification 

Ryan M. Swanson, Justinus A. Satrio, and 

Robert C. Brown

Iowa State University

Alexandru Platon

ConocoPhillips Company

David D. Hsu

National Renewable Energy Laboratory

Prepared under Task No. BB07.7510

Technical Report

NREL/TP-6A20-46587

November 2010
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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 is 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, these results are not indicative of potential performance, but are meant to 

represent the most likely performance given the current state of public knowledge. 
iv 

List of Acronyms 

AGR acid gas removal 

ASU air separation unit 

BTL biomass to liquids 

CFB circulating fluidized bed 

DCFROR discounted cash flow rate of return 

DME dimethyl-ether 

FCI fixed capital investment 

FT Fischer-Tropsch 

GGE gallon of gasoline equivalent 

HRSG heat recovery steam generator 

HT high temperature 

IC indirect costs 

IGCC integrated gasification combined cycle 

IRR internal rate of return 

ISU Iowa State University 

LHV lower heating value 

LT low temperature 

MEA monoethanolamine 

MJ megajoule 

MM million 

MTG methanol to gasoline 

MW megawatt 

Nm3

 normal cubic meter 

NREL National Renewable Energy Laboratory 

PSA pressure swing adsorption 

PV product value 

Sasol South African Coal, Oil, and Gas Corporation 

SPR slurry phase reactor 

SMR steam methane reforming 

SWGS sour water-gas-shift 

TCI total capital investment 

TDIC total direct and indirect cost 

TIC total installed cost 

tpd tons per day 

TPEC total purchased equipment cost 

WGS water-gas-shift 
v 

Executive Summary 

This study compares capital and production costs of two biomass-to-liquid production plants 

based on gasification. The goal is to produce liquid transportation fuels via Fischer-Tropsch 

synthesis with electricity as a co-product. The biorefineries are fed by 2,000 metric tons per day 

of corn stover. The first biorefinery scenario is an oxygen-fed, low-temperature (870°C), nonslagging, fluidized bed gasifier. The second scenario is an oxygen-fed, high-temperature 

(1,300°C), slagging, entrained flow gasifier. Both are followed by catalytic Fischer-Tropsch 

synthesis and hydroprocessing to naphtha-range (gasoline blend stock) and distillate-range 

(diesel blend stock) liquid fractions. (Hydroprocessing is a set of refinery processes that removes 

impurities and breaks down large molecules to fractions suitable for use in commercial 

formulations.) 

Process modeling software (Aspen Plus) is utilized to organize the mass and energy streams and 

cost estimation software is used to generate equipment costs. Economic analysis is performed to 

estimate the capital investment and operating costs. A 20-year discounted cash flow rate of 

return analysis is developed to estimate a fuel product value (PV) at a net present value of zero 

with 10% internal rate of return. All costs are adjusted to the year 2007. The technology is 

limited to commercial technology available for implementation in the next 5–8 years, and as a 

result, the process design is restricted to available rather than projected data. 

Results show that the total capital investment required for nth plant scenarios is $610 million and 

$500 million for high-temperature and low-temperature scenarios, respectively. PV for the hightemperature and low-temperature scenarios is estimated to be $4.30 and $4.80 per gallon of 

gasoline equivalent (GGE), respectively, based on a feedstock cost of $75 per dry short ton. The 

main reason for a difference in PV between the scenarios is because of a higher carbon efficiency 

and subsequent higher fuel yield for the high-temperature scenario. Sensitivity analysis is also 

performed on process and economic parameters. This analysis shows that total capital investment 

and feedstock cost are among the most influential parameters affecting the PV, while least 

influential parameters include per-pass Fischer-Tropsch-reaction-conversion extent, inlet 

feedstock moisture, and catalyst cost. 

In order to estimate the cost of a pioneer plant (first of its kind), an analysis is performed that 

inflates total capital investment and deflates the plant output for the first several years of 

operation. Base case results of this analysis estimate a pioneer plant investment to be $1.4 billion 

and $1.1 billion for high-temperature and low-temperature scenarios, respectively. Resulting PVs 

are estimated to be $7.60/GGE and $8.10/GGE for high-temperature and low-temperature 

pioneer plants, respectively. 
vi 

Table of Contents 

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

Background ................................................................................................................................................. 2

Biorenewable Resources ............................................................................................................2

Gasification ................................................................................................................................2

Reactions ..............................................................................................................................3

Gasifier Types ......................................................................................................................4

Biomass Preprocessing ..............................................................................................................7

Syngas Cleaning.........................................................................................................................9

End-Use Product ......................................................................................................................10

Power Generation ...............................................................................................................10

Synthetic Fuels and Chemicals ..........................................................................................11

Methanol to Gasoline .........................................................................................................11

Fischer-Tropsch .................................................................................................................12

Techno-Economic Analysis .....................................................................................................13

Methodology .............................................................................................................................................. 16

Down-Selection Process ..........................................................................................................16

Preliminary Criteria ...........................................................................................................17

Scenario Selection ..............................................................................................................17

Scenarios Not Selected ......................................................................................................18

Project Assumptions ..........................................................................................................18

Process Description ..................................................................................................................19

High-Temperature Scenario Overview ..............................................................................19

Low-Temperature Scenario Overview ...............................................................................20

Area 100 Preprocessing .....................................................................................................22

Area 200 Gasification ........................................................................................................23

Area 300 Syngas Cleaning .................................................................................................25

Area 400 Fuel Synthesis ....................................................................................................27

Area 500 Hydroprocessing ................................................................................................29

Area 600 Power Generation ...............................................................................................29

Area 700 Air Separation ....................................................................................................30

Methodology for Economic Analysis ......................................................................................30

Methodology for Major Equipment Costs .........................................................................34

Methodology for Sensitivity Analysis ...............................................................................35

Methodology for Pioneer Plant Analysis ...........................................................................36

Results and Discussion ............................................................................................................................ 39

Process Results.........................................................................................................................39

Cost Estimating Results ...........................................................................................................41

Capital and Operating Costs for nth Plant ..........................................................................41

Sensitivity Results for nth Plant ..........................................................................................43

Pioneer Plant Analysis Results ..........................................................................................45

Comparison with Previous Techno-Economic Studies ............................................................46

Summary of nth Plant Scenarios ...............................................................................................48

Conclusions ............................................................................................................................................... 49

References ................................................................................................................................................. 50
vii 

Appendix A. Techno-Economic Model Assumptions ........................................................................... 55

Appendix B. Detailed Costs ..................................................................................................................... 59

Cost Summaries .......................................................................................................................59

Detailed Equipment Lists .........................................................................................................61

Discounted Cash Flow .............................................................................................................67

Appendix C. Scenario Modeling Details ................................................................................................. 71

Property Method ......................................................................................................................71

Stream/Block Nomenclature ....................................................................................................71

Aspen Plus Calculator Block Descriptions ..............................................................................73

Aspen Plus Model Design Specifications ................................................................................84

Detailed Calculations ...............................................................................................................86

Appendix D. Process Flow Diagrams ................................................................................................... 116

High-Temperature Scenario ...................................................................................................117

Low-Temperature Scenario ...................................................................................................127

Appendix E. Stream Data ....................................................................................................................... 138

High-Temperature Scenario ...................................................................................................139

Low-Temperature Scenario ...................................................................................................146
viii 

List of Figures 

Figure 1. Overall process flow diagram for both scenarios ............................................................ 1

Figure 2. Typical energy content of the products of gasification of wood using air varied by 

equivalence ratio [12] ......................................................................................................... 4

Figure 3. Design of fixed-bed (a) updraft and (b) downdraft gasifiers showing reaction zones[13]

............................................................................................................................................. 5

Figure 4. Fluidized bed gasifier designs of (a) and (b) directly heated type and (c) and (d) 

indirectly heated type [16] .................................................................................................. 6

Figure 5. Entrained-flow gasifier [18] ............................................................................................ 7

Figure 6. Schematic of a biomass pretreatment via fast pyrolysis followed by an entrained-flow 

gasifier [17] ......................................................................................................................... 8

Figure 7. Main syngas conversion pathways [32] ........................................................................ 11

Figure 8. Fischer-Tropsch reactor types (a) multi-tubular fixed bed and (b) slurry bed [31] ....... 13

Figure 9. Overall process flow diagram for HT scenario ............................................................. 20

Figure 10. Overall process flow diagram for LT scenario ............................................................ 22

Figure 11. Fischer-Tropsch product distribution as a function of chain growth factor (ࢻ (using 

equation 11 [48] ................................................................................................................ 29

Figure 12. Sensitivity results for HT nth plant scenario ................................................................ 44

Figure 13. Sensitivity results for LT nth plant scenario ................................................................. 44

Figure 14. The effect of plant size on product value (per gallon of gasoline equivalent) for nth

plant scenarios ................................................................................................................... 45

Figure 15. The effect of plant size on total capital investment for nth plant scenarios ................. 45

Figure B-1. Economic analysis summary for HT scenario ........................................................... 59

Figure B-2. Economic analysis summary for LT scenario ........................................................... 60

Figure C-1. Stream nomenclature used in model .......................................................................... 71

Figure C-2. Block nomenclature used in model ........................................................................... 71

Figure C-3. Heat and work stream nomenclature used in model .................................................. 72

Figure C-4. Decision diagram for carbon balance ........................................................................ 80

Figure C-5. Decision diagram for hydrogen balance .................................................................... 81

Figure C-6. Decision diagram for oxygen balance ....................................................................... 82

Figure D-1. Overall plant area process flow diagram for HT scenario ...................................... 117

Figure D-2. Preprocessing area process flow diagram for HT scenario ..................................... 118

Figure D-3. Gasification area process flow diagram for HT scenario ........................................ 119

Figure D-4. Syngas cleaning area process flow diagram for HT scenario ................................. 120

Figure D-5. Acid gas removal area process flow diagram for HT scenario ............................... 121

Figure D-6. Sulfur recovery area process flow diagram for HT scenario .................................. 122

Figure D-7. Fuel synthesis area process flow diagram for HT scenario ..................................... 123

Figure D-8. Hydroprocessing area process flow diagram for HT scenario ................................ 124

Figure D-9. Power generation area process flow diagram for HT scenario ............................... 125

Figure D-10. Air separation unit process flow diagram for HT scenario ................................... 126

Figure D-11. Overall plant area process flow diagram for LT scenario ..................................... 127

Figure D-12. Preprocessing area process flow diagram for LT scenario ................................... 128

Figure D-13. Gasification area process flow diagram for LT scenario ...................................... 129

Figure D-14. Syngas cleaning area process flow diagram for LT scenario ................................ 130

Figure D-15. Acid gas removal area process flow diagram for LT scenario .............................. 131
ix 

Figure D-16. Sulfur recovery process flow diagram for LT scenario ......................................... 132

Figure D-17. Fuel synthesis area process flow diagram for LT scenario ................................... 133

Figure D-18. Syngas conditioning area process flow diagram for LT scenario ......................... 134

Figure D-19. Hydroprocessing area process diagram for LT scenario ....................................... 135

Figure D-20. Power generation area process flow diagram for LT scenario .............................. 136

Figure D-21. Air separation unit process flow diagram for LT scenario .................................... 137

List of Tables 

Table 1. Reactions Occurring within the Reduction Stage of Gasification .................................... 3

Table 2. Previous Techno-Economic Studies of Biomass-Gasification Biofuel Production 

Plants ................................................................................................................................. 15

Table 3. Process Configurations Considered in Down Selection Process .................................... 16

Table 4. Main Assumptions Used in nth Plant Scenarios .............................................................. 18

Table 5. Stover and Char Elemental Composition (wt %) ............................................................ 23

Table 6. Syngas Composition (Mole Basis) Leaving Gasifier for Gasification Scenarios 

Evaluated........................................................................................................................... 25

Table 7. Fischer-Tropsch Gas Cleanliness Requirements [31] ..................................................... 27

Table 8. Hydroprocessing Product Distribution [49] .................................................................... 29

Table 9. Main Economic Assumptions for nth Plant Scenarios .................................................... 30

Table 10. Methodology for Capital Cost Estimation for nth Plant Scenarios ............................... 32

Table 11. Variable Operating Cost Parameters Adjusted to $2007 .............................................. 33

Table 12. Sensitivity Parameters for nth Plant Scenarios .............................................................. 35

Table 13. Pioneer Plant Analysis Parameters and Factors ............................................................ 38

Table 14. Power Generation and Usage ........................................................................................ 39

Table 15. Overall Energy Balance on LHV Basis ........................................................................ 40

Table 16. Overall Carbon Balance ................................................................................................ 41

Table 17. Capital Investment Breakdown for nth Plant Scenarios ................................................ 42

Table 18. Annual Operating Cost Breakdown for nth Plant Scenarios ......................................... 43

Table 19. Catalyst Replacement Costs for Both Scenarios (3-Year Replacement Period) .......... 43

Table 20. Pioneer Plant Analysis Results ..................................................................................... 46

Table 21. Comparison of nth Plant LT Scenario to Tijmensen et al. Study IGT-R Scenario ........ 47

Table 22. Comparison of nth Plant LT Scenario to Larson et al. Study FT-OT-VENT Scenario . 48

Table 23. Main Scenario nth Plant Results .................................................................................... 48

Table B-1. Detailed Equipment List for Areas 100 and 200 of HT Scenario ............................... 61

Table B-2. Detailed Equipment List for Areas 300, 400, and 500 of HT Scenario ...................... 62

Table B-3. Detailed Equipment List for Areas 600 and 700 of HT Scenario ............................... 63

Table B-4. Detailed Equipment List for Areas 100 and 200 of LT Scenario ............................... 64

Table B-5. Detailed Equipment List for Areas 300, 400, and 500 of LT Scenario ...................... 65

Table B-6. Detailed Equipment List for Areas 600 and 700 of LT Scenario ............................... 66

Table B-7. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of HT 

Scenario............................................................................................................................. 67

Table B-8. Discounted Cash Flow Sheet for Years 9-20 of HT Scenario .................................... 68

Table B-9. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of LT Scenario69

Table B-10. Discounted Cash Flow Sheet for Years 9-20 of LT Scenario ................................... 70

Table C-1. Detailed Description of Stream and Block Nomenclature .......................................... 72
x 

Table C-2. Reaction Extent Equations for Each Alkane Hydrocarbon ........................................ 75

Table C-3. Combustion Reactions to Determine Required Oxygen ............................................. 77

Table C-4. Hydroprocessing Product Blend ................................................................................. 84

Table E-1. Overall Plant Stream Data for HT Scenario .............................................................. 139

Table E-2. Preprocessing Area Stream Data for HT Scenario .................................................... 140

Table E-3. Gasification Area Stream Data for HT Scenario ...................................................... 141

Table E-4. Syngas Cleaning Area Stream Data for HT Scenario ............................................... 142

Table E-5. Acid Gas Removal and Sulfur Recovery Areas Stream Data for HT Scenario ........ 143

Table E-6. Fuel Synthesis Area Stream Data for HT Scenario ................................................... 144

Table E-7. Hydroprocessing, Power Generation, and Air Separation Areas Stream Data for 

HT Scenario .................................................................................................................... 145

Table E-8. Overall Plant Stream Data for LT Scenario .............................................................. 146

Table E-9. Preprocessing Area Stream Data for LT Scenario .................................................... 147

Table E-10. Gasification Area Stream Data for LT Scenario ..................................................... 148

Table E-11. Syngas Cleaning Area Stream Data for LT Scenario ............................................. 149

Table E-12. Acid Gas Removal and Sulfur Recovery Areas Stream Data for LT Scenario ...... 150

Table E-13. Fuel Synthesis Area Stream Data for LT Scenario ................................................. 151

Table E-14. Syngas Conditioning Area Stream Data for LT Scenario ....................................... 152

Table E-15. Hydroprocessing, Power Generation, and Air Separation Areas Stream Data for 

LT Scenario ..................................................................................................................... 153
1 

Introduction 

This study investigates the economic feasibility of the thermochemical pathway of gasification of 

biomass to renewable transportation fuels. The objective is to compare capital investment costs 

and production costs for nth plant biorefinery scenarios based on gasification. The selected 

scenarios are high-temperature (slagging) gasification and low-temperature (dry-ash) 

gasification, each followed by Fischer-Tropsch synthesis and hydroprocessing. They are 

designed to produce liquid hydrocarbon fuels from 2,000 dry metric tons (2,205 dry short tons) 

per day of agricultural residue, namely corn stover. Corn stover is chosen as a feedstock in order 

to facilitate comparisons with biochemical and pyrolysis biofuels scenarios [1, 2]. 

The two scenarios were chosen from many options according to the following criteria: 

1. The technology under consideration should be commercially ready in the next 5–8 years. 

2. The size of the biorefinery should be feasible with current agricultural productivity and 

within a realistic feedstock collection area. 

3. In addition, the desired end product should be compatible with the present fuel 

infrastructure, i.e., gasoline and/or diesel. 

The high-temperature gasification scenario is based on a steam/oxygen-fed entrained flow, 

slagging gasifier similar to that described in Frey and Akunuri [3]. The low-temperature 

gasification scenario is based on a pressurized, steam/oxygen-fed fluidized bed gasifier 

developed by the Gas Technology Institute and reported by Bain [4]. The main areas of operation 

are feedstock preprocessing, gasification, syngas cleaning, syngas conditioning/upgrading, fuel 

synthesis, power generation, and air separation (for oxygen production), as shown in Figure 1. 

Process modeling software is utilized to organize the mass and energy streams, and cost 

estimation software is used to generate equipment costs. Economic analysis is performed to 

estimate the capital investment and operating costs. A 20-year discounted cash flow rate of 

return (DCFROR) analysis is developed to estimate a fuel product value (PV) at a net present 

value of zero with 10% internal rate of return. All costs are adjusted to the year 2007. 

Figure 1. Overall process flow diagram for both scenarios 
2 

Background 

Biorenewable Resources 

The world population has long utilized materials that were in close proximity. The nearest 

resource available to the human population is the organic matter in the environment around it. 

This organic matter is present for a limited amount of time due to its decomposable nature. 

Brown [5] defines this material, or biorenewable resources, as organic material of recent 

biological origin. It is a renewable resource if the rate of consumption is equal to the 

regeneration or growth and therefore must be used only if preserving biodiversity [6]. As a 

result, these resources have been important contributors to the world economy, providing 

foodstuffs, transportation, energy, and construction materials, as well as serving many other 

functions. 

Biorenewable resources for generating energy can be classified as woody biomass, energy crops, 

residues, and municipal waste [6]. The first two are primary resources while the remaining are 

secondary resources, meaning that their primary use has already occurred. Woody biomass 

includes logging products and energy crops include short-rotation trees such as poplar and fastgrowing grasses such as switchgrass. Residues can come from logging processing or agricultural 

processing (e.g., corn stover). According to Perlack et al. [7], the energy crop and agricultural 

residue potential in the United States is 1.4 billion annual tons. According to the U.S. 

Department of Energy’s “Roadmap for Agriculture Biomass Feedstock Supply in the U.S.,” 

there is potential for 2 billion annual tons, including municipal waste and biosolids such as 

manure. 

Many end products can be produced from these resources. Aside from the conventional use of 

biomass for human food consumption, livestock feed, and building materials, there are many 

new pathways to provide renewable alternatives to our transportation, infrastructure, and energy. 

Combustion of biomass offers a way to provide heat and power and displaces coal and fuel oil. 

Fermentation of carbohydrates and liquefaction of biomass through fast pyrolysis yield liquid 

products with the potential to displace petrochemicals. Gasification of biomass allows for 

chemical and liquid fuel synthesis, which is the focus of this study. 

Developing an economy that involves biorenewable resources, especially biofuels, has many 

benefits. According to Greene et al. [8], biofuel production has the potential to provide a new 

source of revenue for farmers by generating $5 billion per year. Additionally, toxic and 

greenhouse gas emissions can be reduced by the use of biofuels. In the same study, Greene et al. 

report that 22% of the United States’s total greenhouse gas emissions could be reduced if 

biofuels were developed to replace half of the petroleum consumption. Arguably, the most 

important benefit of biofuel production is the potential for closing the carbon cycle. 

Gasification 

Gasification is a high-temperature and catalytic pathway for producing biofuels. It is defined as 

the partial oxidation of solid, carbonaceous material with air, steam, or oxygen into a flammable 

gas mixture called producer gas or synthesis gas [5]. The synthesis gas contains mostly carbon 

monoxide and hydrogen with various amounts of carbon dioxide, water vapor, and methane. 

Typical volumetric energy content of synthesis gas is 4–18 MJ/Nm3

 [9]. Comparatively, natural 

gas (composed of mostly methane) energy content is 36 MJ/Nm3

 [9]. Much of the energy content 
3 

of the biomass is retained in the gas mixture by partial oxidation rather than full oxidation of the 

biomass, which would result in the release of mostly thermal energy. Historically, gasification of 

coal and wood produced “town gas,” which was subsequently burned in street lamps [10]. 

Additionally, during the World Wars, vehicles were adapted to operate with gasification reactors 

[10]. During this same time period, Germany developed the catalytic synthesis of transportation 

fuels from synthesis gas [11]. The same concept is still in use today by the South African Coal, 

Oil, and Gas Corporation (Sasol) to produce motor fuels and liquid by-products using coal [11]. 

Reactions 

Four stages occur during gasification of carbonaceous material: drying, devolatilization, 

combustion, and reduction [9]. First, the moisture within the material is heated and removed 

through a drying process. Second, continued heating devolatilizes the material where volatile 

matter exits the particle and comes into contact with the oxygen. Third, combustion occurs, 

where carbon dioxide and carbon monoxide are formed from carbon and oxygen. The 

combustion stage is very exothermic and provides enough heat for the last stage, the reduction 

reactions, to occur. The last stage includes water gas reaction, Boudouard reaction, water-gasshift reaction, and methanation reaction (Table 1). As all these stages progress, solid fixed 

carbon remains present. The amount of fixed carbon varies depending on the equivalence ratio. 

Table 1. Reactions Occurring within the Reduction Stage of Gasification 

Name Reaction 

Water Gas ܥ൅ܪଶܱ ՜ ܥܱ ൅ ܪଶ

ܱܥ2 ՜ ଶܱܥ ൅ ܥ Boudouard

Water-Gas-Shift ܥܱ ൅ ܪଶܱ ՜ ܥܱଶ ൅ ܪଶ

Methanation ܥܱ ൅ 3ܪଶ ՜ ܪܥସ ൅ ܪଶܱ

When the equivalence ratio (defined as the actual air/fuel ratio all divided by the stoichiometric 

air/fuel ratio) increases, solid fixed carbon (i.e., char) decreases until enough oxidizer is available 

for complete conversion (Figure 2). This point of complete conversion occurs at an equivalence 

ratio of approximately 0.25. At nearly the same point, the maximum synthesis gas energy content 

(without accounting for sensible energy) is reached. 
4 

Figure 2. Typical energy content of the products of gasification of wood using air varied by 

equivalence ratio [12] 

Gasifier Types 

Three types of reactors are used for gasification: fixed bed, fluidized bed, and entrained flow 

[13]. Fixed beds are fed with biomass from the top of the reactor and form a bed, which gasifies 

as air moves through the bed (Figure 3). As the material releases volatile components, the char 

and ash exit through a grate at the bottom. Typical operating temperature range is 750°–900°C. 

The two main types of fixed-bed gasifiers are updraft and downdraft. The advantage of fixed-bed 

technology is its simplicity, but the technology is limited in scale-up and has low heat mixing 

due to high channeling potential within the reactor [14]. 

0

5

10

15

20

25

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

Enthalpy (MJ/kg biomass, moisture&ash-free)

Equivalence ratio (mol/mol)

total

gas, total

gas, sensible heat

gas, heating value

char, total

char, sensible heat

char, 

heating 

value
5 

 

(a) (b) 

Figure 3. Design of fixed-bed (a) updraft and (b) downdraft gasifiers showing reaction zones [13] 

When the volumetric gas flow is increased through the grate, the fixed bed becomes a fluidized 

bed. Fluidized-bed gasifiers are so named because of the inert bed material that is fluidized by 

oxidizing gas creating turbulence through the bed material (Figure 4). Biomass enters just above 

the top of the bed and mixes with hot, inert material, creating very high heat and mass transfer. 

Operating temperature range is the same as for the fixed bed. Advantages of the fluidized bed 

include flexible feeds, uniform temperature distribution across the bed, and large volumetric flow 

capability [15]. The main types of fluidized-bed gasifiers are circulating fluidized bed (CFB) and 

bubbling fluidized bed. Bubbling-bed gasifiers are directly heated from the combustion reactions 

occurring in the bed. They produce gas and the ash and char fall out the bottom or the side. The 

CFB recycles the char through a cyclone while the product leaves out the top of the cyclone. 

There are also indirectly heated fluidized beds that use a hot material such as sand to provide the 

heat needed for gasification, as shown in Figure 4. Fluidized beds have high carbon conversion 

efficiencies and can scale up easily [14]. 
6 

(a) Bubbling fluidized-bed gasifier (b) Circulating fluidized-bed gasifier 

(c) Indirectly heated gasifier via 

combustor 

(d) Indirectly heated gasifier via heat 

exchange tubes 

Figure 4. Fluidized bed gasifier designs of (a) and (b) directly heated type and (c) and (d) indirectly 

heated type [16] 

Another type of gasifier is the entrained-flow gasifier (Figure 5). Normally operated at elevated 

pressures (up to 50 bar), it requires very fine fuel particles gasified at high temperatures to ensure 

complete gasification during the short residence times in the reactor. The Energy Research 

Centre of the Netherlands has investigated this type of gasification and reported promise with 

biomass, as long as the biomass is pretreated to certain requirements [17]. To keep the residence 

time at approximately the time for a particle to fall the length of the reaction zone, fuel particles 

smaller than 1 mm and high temperatures (1100°–1500°C) are necessary for successful 

operation. 
7 

Figure 5. Entrained-flow gasifier [18] 

Entrained-flow gasification mixes the fuel with a steam/oxygen stream to form a turbulent flow 

within the gasifier. Ash-forming components melt in the gasifier and form a liquid slag on the 

inside wall of the gasifier, effectively protecting the wall itself. The liquid flows down and is 

collected at the bottom. To form the slag, limestone can be added as a fluxing material. For 

herbaceous biomass, such as switchgrass or corn stover, which is high in alkali content, there 

may be sufficient inherent fluxing material present [18]. Advantages of entrained-flow 

gasification are that tar and methane content are negligible and high carbon conversion occurs 

due to more complete gasification of the char. Syngas cleanup is simplified because slag is 

removed at the bottom of the gasifier, negating the need for cyclones and tar removal [19]. The 

disadvantages are that very high temperatures need to be maintained and the design and 

operation is more complex. An entrained-flow gasifier co-firing up to 25% biomass with coal has 

been developed by Shell in Buggenum, Netherlands. Another gasifier developed by Future 

Energy in Freiburg, Germany, uses waste oil and sludges. Both are operating at commercial scale 

[17]. 

Biomass Preprocessing 

A degree of biomass processing is required before gasification can occur. Most gasifiers require 

smaller size feedstock than is typically collected during harvest. Therefore, a significant degree 

of size reduction is necessary. A typical setup for size reduction is using a two-step process in 

which a chipper accomplishes the primary reduction and a hammer mill is used for the secondary 

reduction [20]. In addition, the maximum moisture content for biomass gasification is 20%–30% 
8 

(wet basis), and for normal operation it is less than 15% (wet basis) [9]. Therefore, a drying 

process is required to prepare the feedstock for gasification. 

The main benefit of drying biomass is to avoid using energy within the gasifier to heat and dry 

the feedstock [21]. Drier biomass also makes for more stable temperature control within the 

gasifier. Rotary dryers typically operate utilizing hot flue gas from a downstream process as the 

drying medium. They have high capacity but require long residence times. In addition, rotary 

dryers have a high fire hazard when using flue gas [21]. To avoid using flue gas, rotary dryers 

can use superheated steam, essentially an inert gas, when a combined-cycle heat and power 

system is used downstream. That system has significant steam available for use because of the 

steam produced in the steam cycle. An advantage of using steam for drying is better heat transfer 

and therefore shorter residence time. 

Pretreatment options for entrained-flow gasification include torrefaction followed by grinding to 

0.1-mm particles, grinding to 1-mm particles, pyrolysis to produce bio-oil/char slurry (bioslurry), 

and initial fluidized-bed gasification of larger particles coupled to an entrained-flow gasifier. 

Torrefaction, essentially an oxygen-free roasting process, causes the biomass particles to be 

brittle, which makes for easy grinding but releases up to 15% of the energy in the biomass via 

volatile compounds [17]. The fluidized/entrained coupled option is attractive because of an 

overall energy efficiency of 80%–85%, but it is expensive because two gasifiers are used in 

series. 

The bioslurry option is illustrated in Figure 6. Basically, a flash-pyrolysis process yields bio-oil 

and char and is followed by a slagging, entrained-flow gasifier. Because this process utilizes an 

entrained-flow gasifier, the feed must be pressurized. Fortunately, the pyrolysis slurry, already in 

an emulsified liquid state, can be pressurized easily. Technology for slurry feeding is state of the 

art as a result of experience with coal slurries [17]. The bioslurry still contains 90% of the energy 

contained in the original biomass [22]. Another advantage is that no inert gas is needed for solids 

pressurization, avoiding dilution of the feed, which would dilute the syngas. In the search for 

cost-effective methods for production of syngas, this option has potential, but it isn’t as 

developed as other technologies such as fluidized-bed gasification. The biggest challenge is 

constructing and operating a large-scale pyrolysis process, because large-scale systems have not 

been demonstrated [17]. 

Figure 6. Schematic of a biomass pretreatment via fast pyrolysis followed by an entrained-flow 

gasifier [17] 
9 

Syngas Cleaning 

Because the raw syngas leaving the gasifier contains particulates, tars, alkali compounds, sulfur 

compounds, nitrogen compounds, and other contaminants, those components must be removed 

or reduced significantly. Particulates and tars have the potential for clogging downstream 

processes. Sulfur and nitrogen have the potential to poison downstream processes, especially 

catalysts used in fuel synthesis applications. Moreover, another motivation for cleaning syngas is 

meeting environmental emissions limits. 

The syngas must be cooled before conventional gas cleanup is utilized. This can be 

accomplished in two ways: direct quench by injection of water and indirect quench via a heat 

exchanger. Direct quench is less expensive but dilutes the syngas. Direct quenching can also be 

used to clean up the gas by removing alkali species, particulates, and tars [23]. 

Particulates are defined as inorganic mineral material, ash, and unconverted biomass or char 

[24]. In addition, bed material from the gasifier is included in the particulates. Switchgrass 

feedstock typically has 10% inorganic material in the form of minerals. Many gasifiers operate 

with a 98%–99% carbon conversion efficiency where 1%–2% of the solid carbon is in the form 

of char [24]. 

Particulates are primarily removed by physical systems such as cyclones, in which the heavy 

particles fall down the center while the gases rise up and out of the cyclone. The initial step for 

particulate removal is usually a cyclone. Importantly, particulates should be removed before the 

gas is cooled down for cold gas cleaning. If removed after gas cooling, tars can condense onto 

the particulates and potentially plug equipment. Alternatively, barrier filters, which operate 

above tar condensation temperatures, use metal or ceramic screens or filters to remove 

particulates while allowing the gas to remain hot. Barrier filters, however, have presented 

problems in sintering and breaking [24]. 

Even more critical to downstream syngas applications is tar removal. Tars are defined as higherweight organics and oxygenated aromatics heavier than benzene (78 g/mol). They are produced 

from volatized material after polymerization [24]. A review by Milne et al. [25] of tars produced 

during gasification covers different removal methods. One method uses physical removal via wet 

gas scrubbing of tars in a scrubbing tower for the “heavy tars” followed by a venturi scrubber for 

lighter tars. This setup is similar to the direct quench cooling mentioned previously, as cooling 

occurs as well. Tar concentration is reported to be lower than 10 ppm by volume at the exit of 

this setup [23]. The disadvantage of this setup is that wastewater treatment is required to dispose 

of the tar and can be expensive. 

The other method for tar removal is catalytic or thermal conversion to non-condensable gas. This 

is also known as hot gas cleaning, as it occurs at temperatures at or above gasification 

temperatures. Catalytic conversion can occur at temperatures as low as 800°C, and thermal 

conversion can occur at temperatures up to 1200°C. The energy required for thermal tar cracking 

may not be cost-competitive because of the temperature rise required from the gasification 

temperature to crack the high refractory tars [24]. 
10 

Alkali compounds such as calcium oxide and potassium oxide are present in biomass. When 

gasified, both either vaporize or concentrate in the ash. Condensation of these compounds begins 

at 650°C, and they can deposit on cool surfaces causing equipment clogging, equipment 

corrosion, and catalyst deactivation [26]. According to Stevens [26], research on alkali 

adsorption filters using bauxite has been promising but not demonstrated on a large scale. 

Stevens concludes that the best current method for alkali removal is using proven syngas cooling 

followed by wet scrubbing, where the addition of water cools the syngas and physically removes 

small particles and liquid droplets. 

Wet scrubbing also removes ammonia that forms during gasification from the nitrogen in the 

biomass. Without proper removal, ammonia can deactivate catalysts as well. Complete ammonia 

removal can be accomplished through wet scrubbing [27]. For gasifiers coupled to a catalytic or 

thermal tar reformer, most of the ammonia can be reformed to hydrogen and nitrogen [27]. 

Sulfur in the biomass mostly forms hydrogen sulfide (H2S) with small amounts of carbonyl 

sulfide (COS). Hydrogen sulfide is removed by three main ways: chemical solvents, physical 

solvents, and catalytic sorbents. For chemical removal, amine-based solvents are typically 

utilized to chemically bond with H2S. Physical removal takes advantage of the high solubility of 

H2S using an organic solvent. Typical setups of both chemical and physical removal involve an 

absorber unit followed by a solvent regenerator unit, known as a stripper. Operation usually 

occurs at temperatures lower than 100°C and medium to high pressures (150–500 psi) [27]. 

Sulfur leaving these two systems is around 1–4 ppm and can require further removal, especially 

for fuel synthesis. In that case, a syngas polishing step using a fixed-bed zinc oxide-activated 

carbon catalyst removes H2S and COS to achieve the parts per billion levels necessary for fuel 

synthesis. Halides, present in trace amounts in the biomass, can also be removed with the zinc 

oxide catalyst [27]. The water streams with the scrubbed-out impurities are sent to wastewater 

treatment. 

End-Use Product 

After syngas is cleaned of particulates, impurities, and contaminants, there is sufficient energy 

content for producing a higher valued product. There are three main large-scale biomass 

gasification pathways that have been researched and suggested for producing higher valued 

products: power generation, liquid-fuel synthesis, and chemical synthesis. According to Wender 

[28], the three largest commercial uses for syngas are ammonia production from hydrogen, 

methanol synthesis, and hydrocarbon synthesis via the Fischer-Tropsch process. 

Power Generation 

Power is most effectively generated using gasification by combusting the syngas in a gas turbine 

to provide mechanical work for a generator. Steam can then be generated by recovering heat 

from the hot syngas, and the steam in turn is used for mechanical work via a steam turbine. This 

gasifier plus gas turbine and steam turbine setup is known as integrated gasification combined 

cycle (IGCC) power generation. The level of particulates, alkali metals, and tar can decrease the 

performance of the gas turbine. Consonni and Larson [29] found that particulates can cause 

turbine blade erosion; 99% of 10-micron or smaller particles should be removed. In addition, 

they also report that alkali metals corrode the turbine blades and tars condense on the turbine 

blades, both hindering operation and escalating turbine failure. Fortunately, nearly all alkali 

metals and tars can be removed using proven wet scrubbing techniques. 
11 

Using the IGCC approach to generate power, Bridgwater et al. [30] and Craig and Mann [23] 

expect biomass to produce power with efficiencies in the range of 35%–40% with large scale 

systems (greater than 100 MW net output) at the high end of the range. Moreover, Craig and 

Mann suggest that future advanced turbine systems could reach 50% biomass-to-power 

efficiency. 

Synthetic Fuels and Chemicals 

Instead of converting the energy content of the syngas to power, the energy content can be 

condensed into a liquid energy carrier, or fuel. The conversion of syngas to fuels can only occur 

in the presence of proper catalysts [31]. The catalytic reactions basically build up the small 

molecules in the syngas (i.e., carbon monoxide and hydrogen) into larger compounds that are 

more easily stored and transported. A summary of many catalytic pathways to fuels and 

chemicals is shown in Figure 7. In most catalytic synthesis reactions, syngas cleanliness 

requirements are very high. Most impurities and contaminants are removed to low parts-permillion and even parts-per-billion concentrations. This means that significant cost must be 

directed toward syngas cleaning. 

Figure 7. Main syngas conversion pathways [32] 

Methanol to Gasoline 

Methanol is one of the top chemicals produced in the world [32]. Most commercially produced 

methanol is synthesized via steam methane reforming and autothermal reforming. The synthesis 

of methanol from syngas is highly exothermic (equation 1). The reaction occurs over a 
12 

Cu/ZnO/Al2O3 catalyst at temperatures of 220°–275°C and pressures of 50–100 bar with a 

catalyst lifetime of 2–5 years [31]. Wender [28] reports syngas-to-methanol conversion 

efficiency can reach 99% with recycle, but per-pass efficiency is only about 25%. 

(1 .eqn (ܪܱଷܪܥ ՜ ଶܪ2 ൅ ܱܥ

Although methanol can be used directly as a liquid fuel, it can also be converted into the 

conventional transportation fuel range. This process is known as the methanol-to-gasoline 

(MTG) process and was developed by the Mobil Oil Corporation [31]. In that process, methanol 

is heated to 300°C and dehydrated over alumina catalyst at 27 atm, yielding methanol, dimethyl 

ether (DME), and water. The exiting mixture reacts with a zeolite catalyst at 350°C and 20 atm 

to produce 56% water and 44% hydrocarbons by weight. Of the hydrocarbon product, 85% is in 

the gasoline range, with 40% of the gasoline-range product being aromatic. However, limitations 

on the aromatic content of gasoline have been proposed in legislation [31]. The thermal 

efficiency of the MTG process is 70% [11]. The overall MTG process usually includes multiple 

MTG reactors in parallel to perform periodic catalyst regeneration by burning off coke deposits 

[11]. Mobil operated a commercial plant producing 14,500 barrels per day in New Zealand 

during the 1980s [32]. Alternatively, the reaction process could be stopped directly after the 

methanol synthesis to focus on producing DME, because DME can be used as a diesel fuel 

because it has a high cetane number. DME is formed from the dehydration reaction of methanol 

over an acid catalyst γ-alumina. Per-pass efficiency can be as high as 50%. Overall, syngas-toDME production efficiency is higher than that for syngas to methanol [31]. However, DME is in 

gaseous form at atmospheric conditions and needs to be pressurized for use in diesel engines 

[33]. Therefore, engine modification is required; this is the main disadvantage for DME use as 

transportation fuel. 

Fischer-Tropsch 

Fischer-Tropsch (FT) catalytic synthesis is a highly exothermic reaction producing a wide 

variety of alkanes (equation 2). 

ܥܱ ൅ 2.1ܪଶ ՜ െሺܪܥଶሻ െ ൅ܪଶܱ (eqn. 2)

For gasoline-range products, higher temperatures (300°–350°C) and iron catalysts are typically 

used. For diesel-range and wax products, lower temperatures (200°–240°C) and cobalt catalysts 

are typically used [34]. Operating pressures are in the range of 10–40 bar. Product distribution 

can be estimated using the Anderson-Schulz-Flory chain-growth-probability model, in which 

longer hydrocarbon chains form as the temperature decreases. At high temperatures, selectivity 

favors methane and light gases. This is a disadvantage if liquid fuel production is the focus. At 

low temperatures, selectivity favors long-carbon-chain wax products requiring further 

hydrocracking to the diesel range in a separate unit, which adds more construction cost but is 

necessary for liquid fuel production. 

Because of the highly exothermic reaction, the heat must be removed or the catalyst can be 

deactivated. Two main types of reactors have been designed: a fixed-bed tubular reactor and a 
13 

slurry-phase reactor (Figure 8). Heat removal is crucial to the process and has been the focus of 

reactor design development [31]. The fixed-bed reactor has many catalyst tubes with heat 

removal achieved by steam generation on the outside of the tubes [35]. The fixed-bed reactor is 

simple to operate and is well suited for wax production due to simple liquid/wax removal. 

However, it is more expensive to build because of the many tubes and has a high pressure drop 

across the reactor [36]. The slurry-phase reactor (SPR) operates by suspending catalyst in a 

liquid and bubbling the syngas through from the bottom. Disadvantages of the SPR are more 

complex operation and difficult wax removal. However, SPRs cost approximately 40% less to 

build than fixed-bed reactors [36]. 

FT diesel is very low in sulfur, low in aromatic content, and has a high cetane number, making it 

very attractive as a conventional fuel alternative. Emissions across the board decrease when 

using FT diesel. A South Africa-based company, Sasol, has been producing transportation fuel 

since 1955 using the FT process and supplies 41% of South Africa’s transportation fuel 

requirements [31]. 

(a) (b) 

Figure 8. Fischer-Tropsch reactor types (a) multi-tubular fixed bed and (b) slurry bed [31] 

Techno-Economic Analysis 

For biofuels technologies to be utilized in commercial applications, the economic feasibility 

must be determined. A feasibility analysis is also called a techno-economic analysis, in which the 

technical aspects of a project are coupled to the economic aspects. First, the basic theoretical 

configuration is developed and a mass and energy balance is performed. Second, cost estimation 
14 

allows the investment and production cost of a biorefinery to be determined. With rising interest 

in biorenewable resources, many techno-economic studies have been performed on power 

generation and biofuel scenarios. These studies assist in understanding how the physical process 

relates to the cost of producing renewable alternatives. Accuracy of results from these studies is 

usually ±30% of the actual cost [5]. 

Previous studies of gasification-based biomass-to-liquid production plants estimate the cost of 

transportation fuels to range from $12/GJ to $16/GJ ($1.60–$2.00 per gallon of gasoline 

equivalent) [16, 37-40]. The same studies estimate total capital investment in the range of $191 

million for 2,000 dry metric ton per day (tpd) input [39] to $541 million for 4,500 dry metric tpd 

input [38]. 

A 1,650 dry metric tpd biomass-to-methanol plant based on gasification, with a production cost 

of $15/GJ ($0.90/gal of methanol), is reported by Williams et al. [16] in 1991$ for $45/dry 

metric ton of biomass. Williams et al. also shows the production cost of methanol-derived natural 

gas to be $10/GJ ($0.60/gal of methanol). However, that study concludes that if a carbon tax 

system was developed for lifecycle carbon emissions, then renewable methanol could become 

competitive with natural-gas-derived methanol at a tax of approximately $90 per metric ton of 

carbon. A more recent study by Larson et al. of switchgrass-to-hydrocarbons production in 2009 

reports a production cost of $15.3/GJ ($1.90/gal of gasoline) in 2003$ for a 4,540 dry metric tpd 

(5,000 dry short tpd) plant based on gasification [38]. 

Table 2 compares four biofuel production studies based on gasification. A range of cost year, 

plant size, and feedstock cost show the diversity of characteristics and assumptions that technoeconomic studies use. In addition, resulting capital investment costs of the studies have a large 

range. For example, the capital investment costs of the Phillips et al. [39] and Tijmensen et al. 

[40] studies are $191 million and $387 million, respectively, at similar plant sizes. Reasons for 

such a significant difference are choice of technologies and level of technology development. 

The Phillips et al. study is a target study, meaning that it assumes future technology 

improvement and therefore lower costs. Direct comparison is difficult because of the varying 

assumptions used by each study. 

 
15 

Table 2. Previous Techno-Economic Studies of Biomass-Gasification Biofuel Production Plants 

Williams et 

al. [16]

Phillips et al. 

[39]

Tijmensen et 

al. [40]

Larson et 

al. [38]

Cost Year 1991 2005 2000 2003 

Plant Size (dry metric 

ton per day) 1,650 2,000 1,741 4,540 

Feedstock Generic 

biomass Poplar Poplar Switchgrass 

Fuel Output Methanol Ethanol FT liquids Diesel, 

gasoline 

Feedstock Cost 

($/dry short ton) 41 35 33 46 

Capital Investment 

($MM) N/A 191 387 541 

Product Value ($/GJ) 15 12 16 15 

Product Value 

($/GGE) 1.90 1.60 2.00 1.85 
16 

Methodology 

The following steps were undertaken to perform the analysis in this study: 

• Collect performance information on relevant technologies for systems under evaluation 

• Perform down-selection process with developed criteria to identify most appropriate 

scenarios 

• Design process models using Aspen Plus process engineering software 

• Size and cost equipment using Aspen Icarus Process Evaluator software, literature 

references, and experimental data 

• Determine capital investments and perform discounted cash flow analysis 

• Perform sensitivity analysis on process and economic parameters 

• Perform pioneer plant cost growth and performance analysis. 

 

Down-Selection Process 

A number of process configurations for the gasification-based, biomass-to-liquids (BTL) route 

were initially considered. These configurations are listed in Table 3 and discussed in the 

following sections. 

Table 3. Process Configurations Considered in Down Selection Process 

Gasifier Block 

Entrained flow, slagging gasifier 

Fluid bed, dry ash gasifier 

Transport gasifier, dry ash (e.g., Kellog, Brown, and Root) 

Indirect gasifier, dry ash (e.g., Battelle-Columbus Labs) 

Syngas Cleaning 

Water scrubbing 

Catalytic tar conversion/reduction 

Thermal tar conversion/reduction 

Amine-based acid gas removal 

Physical sorbent-based acid gas removal (e.g., Selexol, Rectisol) 

Fuel Synthesis 

Fischer-Tropsch 

Mixed alcohols 

Methanol to gasoline (MTG) 

Dimethyl ether 

Syngas fermentation 
17 

Preliminary Criteria 

The initial technology configuration options are reviewed and screened in accordance with three 

criteria. 

First, the technology under consideration should be commercially ready in the next 5–8 years 

and preferably with high technology development. High technology development increases the 

likelihood that a configuration will perform at the scale defined in this study. For example, coal 

gasification has been demonstrated commercially at large scales [11]. While similar scale 

biomass gasifiers have not been proven commercially, the technology development of coal is 

assumed to apply for biomass in 5–8 years. 

Second, the size of biorefinery should be feasible with typical agricultural productivity and 

within a realistic collection area. For example, if one-third of total land use surrounding the 

biorefinery is for stover collection and each acre provides conservatively 1 short dry ton per year, 

then the required collection radius is 35 miles and the amount of biomass transported to the 

biorefinery is approximately 2,300 short tons (2,090 metric tons) per day. The collection area 

with a 35-mile radius is assumed to be realistic. In addition, previous studies by Tijmensen et al., 

Phillips et al., and Lau et al. have used similar plant sizes [39-41]. 

Third, the desired product should be compatible with the present transportation fuel 

infrastructure, i.e., gasoline and diesel-range hydrocarbons. 

Scenario Selection 

For the gasification area, two gasifiers are selected for modeling. First, an entrained-flow, 

slagging gasifier is chosen because of its commercial application with coal (GE, Siemens, Shell, 

and ConocoPhillips) and its potential for use with biomass. Moreover, process modeling of this 

gasifier is simple because it can be closely approximated at thermodynamic equilibrium [3]. 

Second, a fluidized-bed, dry-ash gasifier is chosen because of experience at the Gas Technology 

Institute and because of data availability. Also, a report by Bain [4] at the National Renewable 

Energy Laboratory contains collected and analyzed data for fluidized-bed gasification. In 

addition, Iowa State University is currently operating an atmospheric-pressure fluidized-bed 

gasifier as either air or oxygen/steam fed. 

The syngas cleaning area is chosen to include configurations that have less technological 

complexity than previous studies. Phillips et al. [39] and Larson et al. [38] both employ an 

external catalytic tar-reforming process for dry-ash gasification. Because of low technological 

development in tar conversion and its inherent complexity, a direct-contact syngas quenching 

and scrubbing system is instead chosen for this study. In the case of the slagging gasifier, high 

temperatures inhibit tar formation, yet quenching and particulate and ammonia removal are still 

required. An amine-based chemical absorber/stripper configuration is chosen for removal of 

hydrogen sulfide and carbon dioxide. This configuration is chosen due to data availability, as 

compared to proprietary physical gas cleaning process such as Rectisol and Selexol. 

Two fuel synthesis configurations under consideration produce liquid hydrocarbons: FischerTropsch (FT) synthesis and MTG. FT synthesis has been proven in operation at commercial 

scale for many years by Sasol [11]. Because of more accessible data and long industrial 

experience, FT synthesis is the only fuel synthesis option chosen. Because of this selection, a 
18 

post-synthesis fuel upgrading area is also necessary, as FT products need to be separated and 

hydroprocessed. 

Scenarios Not Selected 

The indirect, dry-ash gasifier and the mixed alcohol synthesis configurations are not considered 

due to previous work by Phillips et al. [39]. The transport gasifier design, though a promising 

technology, is not considered because of reactor complexity, unproven commercial-scale 

operation, and lack of public domain data. Tar conversion via external thermal or catalytic 

cracking is not considered due to lack of public domain data and commercial scale experience. 

Acid gas removal using proprietary technology (e.g., Rectisol or Selexol) is not considered 

because of a lack of public operational data. MTG, including methanol synthesis, is not 

considered because of time constraints and limited operational data. DME and syngas 

fermentation are not considered because of limited commercial scale experience and 

incompatibility with present fuel infrastructure. 

Project Assumptions 

The main project assumptions for process and economic analysis are listed in Table 4. A more 

extensive list can be found in Appendix A. The process design is assumed to incorporate an “nth

plant” level of implementation experience. This is a significant assumption for a process concept 

that has yet to undergo detailed engineering development and see its first commercial 

application. The design and operation of an nth plant is likely to diverge from the design and 

operation considered in this study. However, the nth plant evaluation is chosen to provide an 

analysis similar to studies completed by other groups, including NREL and Pacific Northwest 

National Laboratory. 

Table 4. Main Assumptions Used in nth Plant Scenarios 

Main Assumptions 

The plant is modeled as nth plant 

Plant capacity is 2,000 dry metric ton/day 

Feedstock is corn stover at 25% moisture 

Feedstock ash content at 6% 

Feedstock is purchased at plant gate for $75/dry short ton 

All financial values are adjusted to 2007 cost year 

Plant is 100% equity financed 

Fuel PV is evaluated at 10% internal rate of return 

Plant initiates operation in 5–8 year time frame 

Plant life is 20 years 

Plant availability is 310 days per year (85%) 
19 

Process Description 

High-Temperature Scenario Overview 

The high-temperature (HT) or slagging scenario is a 2,000 dry metric ton (2,205 dry short ton) 

per day corn stover-fed gasification biorefinery that produces naphtha-range and distillate-range 

liquid fractions to be used as blend stock for gasoline and diesel, respectively, as well as 

electricity for export. It is based on pressurized, oxygen-blown, entrained-flow gasification. The 

HT scenario is an nth plant design, meaning significant design, engineering, and operating 

experience has been achieved. 

Figure 9 shows seven main areas of operation for the HT scenario. Feedstock preprocessing 

(Area 100) is where the stover is chopped, dried, and ground to 1-mm size and 10% moisture. 

Gasification (Area 200) contains the stover pressurization for solids feeding, gasification, and 

slag removal. Synthesis gas cleaning (Area 300) contains cold gas cleaning technologies where 

the syngas is quenched and scrubbed of particulates, ammonia, hydrogen sulfide, and carbon 

dioxide. Area 300 also contains the water-gas-shift reaction, which occurs before the hydrogen 

sulfide and carbon dioxide removal. This adjusts the ratio of hydrogen to carbon monoxide for 

optimal fuel synthesis. Fuel synthesis (Area 400) contains syngas boost pressurization, 

contaminant polishing via zinc oxide guard beds, Fischer-Tropsch reactor, and hydrocarbon 

gas/liquid separation. Hydroprocessing (Area 500) produces the final fuel blend and is treated as 

a black box utilizing published data. Power generation (Area 600) contains gas and steam 

turbines along with a heat recovery steam generator. Area 700 contains the air separation unit 

(ASU) where oxygen is separated from air and pressurized for use in the gasifier. For cost 

analysis uses only, a balance-of-plant area includes the cooling tower area, cooling water system, 

waste solids and liquids handling area, and feed water system. Detailed process flow diagrams 

can be found in Appendix D and detailed stream data can be found in Appendix E. 
20 

Dryer

Acid Gas

Removal

Water,

Solids

Fischer

Tropsch Hydroprocessing

Naphtha, Distillate

to storage

Fuel Gas

to A600

Slag

Grinder Entrained

Flow

Gasifier

Lockhopper

Water

Sour Water

Gas Shift

LO-CAT

Sulfur

Cake

Compressor

Pressure Swing

Adsorption

Hydrogen

Power

Combustion

Flue Gas

Steam

Syngas

A700

A100

A200 A300 A400 A500

A600 Chopper

Power Generation

Steam

Stover

Air

Condensate

Nitrogen

Oxygen

Dry Stover

Syngas

Carbon Dioxide

1

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

21

22

24

23

25

26 27

28

29

33

Oil/Water

Separation

Water 34

31

32

30

Direct Quench

Stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17

T (°C) 25 25 90 90 16 32 149 120 200 120 220 249 1300 50 203 203 190

P (bar) 1.01 1.01 1.01 1.01 1.1 1.01 28 1.98 1.98 1.98 1.03 28 26.62 26.62 26.62 25.93 10

m (Mg/day) 2667 2667 2222 2222 2189 2903 744 444 4000 4000 271 180 3823 114 4000 3867 550

Stream 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34

T (°C) 240 40 40 50 50 62 76 30 200 42 45 45 45 35 35 35 37

P (bar) 24.8 24.82 24.82 3.45 1.93 22.75 26 1.01 24.96 23.6 23.58 23.58 23.58 22.2 22.2 22.2 1.03

m (Mg/day) 4417 1502 2965 1755 3.1 3570 3570 4.38 1404 1068 2308 156 33 427 641 52 378

Figure 9. Overall process flow diagram for HT scenario (parallelograms enclosing numbers in the 

diagram designate individual process streams, which are detailed in the accompanying table) 

Recycle streams are utilized to provide better syngas-to-FT-products conversion. Unconverted 

syngas in the fuel synthesis area is recycled to the syngas cleaning area to remove carbon dioxide 

and allow for further conversion in the Fischer-Tropsch reactor. A small portion of unconverted 

syngas is sent to a steam boiler to raise steam required for drying the biomass. The balance of 

unconverted syngas is combusted in a gas turbine and waste heat is recovered in a steam 

generator for steam turbine power. Power generated is used throughout the plant and excess is 

sold. 

Some of the largest consumers of power are the ASU and hydroprocessing area at 11.6 MW and 

2.2 MW, respectively. Another consumer of power is the hammer mill for grinding the dried 

biomass in Area 100; it requires 3.0 MW. The amine/water solution recirculation pump in Area 

300 requires approximately 0.9 MW. Syngas compressors throughout the plant require a 

significant amount of power as well. Gross plant power production is 48.6 MW and net 

electricity for export is 13.8 MW. 

Low-Temperature Scenario Overview 

The low-temperature (LT) or dry-ash scenario is a 2,000 dry metric ton (2,205 dry short ton) per 

day corn stover-fed gasification biorefinery that produces naphtha-range and distillate-range 
21 

liquid fractions to be used as blend stock for gasoline and diesel, respectively, as well as 

electricity for export. It is based on a pressurized, oxygen/steam-blown fluidized-bed gasifier 

developed by Gas Technology Institute. The LT scenario is an nth plant design, meaning 

significant design, engineering, and operating experience has been achieved. 

Figure 10 shows the seven main areas of operation for the LT scenario. Feedstock preprocessing 

(Area 100) is where the stover is chopped, dried, and ground to 6-mm size and 10% moisture. 

Gasification (Area 200) contains the stover pressurization for solids feeding, gasification, and 

char and ash removal. Synthesis gas cleaning (Area 300) contains cold gas cleaning technologies 

where the syngas is quenched and scrubbed of particulates, ammonia, hydrogen sulfide, and 

carbon dioxide. Fuel synthesis (Area 400) contains syngas boost pressurization, contaminant 

polishing via zinc oxide beds, Fischer-Tropsch reactor, and hydrocarbon gas/liquid separation. 

Also included within Area 400 is the steam methane reformer (SMR) to reduce methane content 

and water-gas-shift (WGS) to adjust the ratio of hydrogen and carbon monoxide. 

Hydroprocessing (Area 500) produces the final fuel blend and is treated as a black box utilizing 

published data. Power generation (Area 600) contains gas and steam turbines along with a heat 

recovery steam generator. Area 700 contains the ASU that separates oxygen from air and 

pressurizes it for use in the gasifier. Detailed process flow diagrams can be found in Appendix D 

and detailed stream data can be found in Appendix E. 
22 

Figure 10. Overall process flow diagram for LT scenario (parallelograms enclosing numbers in the 

diagram designate individual process streams, which are detailed in the accompanying table) 

Recycle streams are utilized to provide better FT products conversion. Unconverted syngas in the 

fuel synthesis area is recycled to the syngas cleaning area to remove carbon dioxide and allow 

for further conversion in the Fischer-Tropsch reactor. The balance of unconverted syngas is 

combusted in a gas turbine and waste heat is recovered in a steam generator for steam turbine 

power. Power generated is used throughout the plant and excess is sold. Unconverted carbon 

within the gasifier in the form of char is collected and combusted in a furnace to produce heat, 

thereby generating steam for the drying of the biomass. 

Some of the largest consumers of power are the ASU and hydroprocessing area at 9.1 MW and 

1.7 MW, respectively. Another consumer of power is the hammer mill for grinding the dried 

biomass in Area 100; it requires 1.1 MW. The amine/water solution recirculation pump in Area 

300 requires approximately 0.7 MW. Syngas compressors throughout the plant require a 

significant amount of power as well. Gross plant power production is 40.7 MW and net 

electricity for export is 16.3 MW. 

Area 100 Preprocessing 

The preprocessing area contains all the unit operations required for preparing the biomass for 

feeding into the gasifier. Biomass enters the plant gate at 25 wt% moisture on wet basis in bales. 

The corn stover composition is shown below in Table 5. Ash content is assumed to be 6% by 

Stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17

T (°C) 25 25 90 90 16 32 149 120 200 120 200 243 870 870 1500 40 40

P (bar) 1.01 1.01 1.01 1.01 1.10 1.01 28.00 1.98 1.98 1.98 1.00 28.00 27.55 27.55 1.00 26.17 26.17

m (Mg/day) 2667 2667 2222 2222 1744 2313 569 444 4000 4000 1471 180 2930 215 119 6000 2198

Stream 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33

T (°C) 40 50 50 62 78 870 304 60 200 32 32 32 35 35 50 50

P (bar) 26.17 3.45 2.10 22.89 26.50 25.81 25.12 1.01 25.00 22.89 22.89 22.89 22.89 22.89 22.89 22.89

m (Mg/day) 1380 1593 2.3 2709 2709 3707 3708 3.4 560 1463 2073 168 331 1132 41 293
23 

weight. The composition of char formed in the gasifier is also shown in Table 5. Forklifts 

transport the bales to conveyors where the stover is separated from any metal in a magnetic 

separator. The first modeled operational area is a primary biomass chopper to complete the initial 

size reduction step and prepare stover for drying. 

Table 5. Stover and Char Elemental Composition (wt %) 

The next area of operation is the direct-contact steam drying, which is modeled as a rotary steam 

dryer with exiting biomass moisture of 10% on wet basis. For steam dryers, Amos [21] suggests 

a 9:1 ratio of steam to evaporated moisture. Therefore, 4,000 metric tpd steam is utilized in a 

loop and heated to 200°C from the hot combustion flue gases exiting the syngas- or char-fired 

combustor in Area 200. Steam mixes with 25°C biomass and enters the dryer. At the exit, steam 

at 120°C returns to the combustor for reheating, and dried biomass exits at 90°C and is conveyed 

to the grinding area. 

The grinding area has the same configuration as the chopping area except that the grinder 

requires significantly more power because of the larger size reduction. The grinder reduces the 

size of the biomass to 1-mm and 6-mm particles for the HT and LT scenarios, respectively. The 

power requirements of the grinder for the HT and LT scenarios are 3,000 kW and 1,100 kW, 

respectively. Energy requirements for grinding are determined using the correlations for specific 

energy (kWh per short ton), which is adapted from Mani et al. [42]. 

Area 200 Gasification 

The gasification area of the plant produces synthesis gas using pressurized gasifiers. Also in this 

area, slag, char, and ash are removed. This area also includes lockhoppers for biomass 

pressurization and a fired combustor that provides heat to raise steam for drying the stover. 

Dried and ground stover enters the area and is immediately conveyed to a lockhopper system for 

pressurized feeding. Carbon dioxide from the syngas cleaning area is used as pressurization gas. 

According to Lau et al. [41], a lockhopper system is the best setup for pressurized feeding of 

solids, despite higher operating costs due to high inert gas usage. A proven track record with 

biomass is the main reason for their recommendation. The power requirement of a lockhopper 

system using biomass is 0.082 kW/metric tpd, resulting in a 180-kW system. Higman and van 

der Burgt [43] report inert gas usage as 0.09 kg/kg for 25-bar applications. This results in a 180 

metric ton per day carbon dioxide addition into the hopper. It is assumed that only 5% of the 

inert gas leaks into the gasifier while the rest is vented by the lockhopper. 

Element Stover Char 

Ash 6.00 0 

Carbon 47.28 68.05 

Hydrogen 5.06 3.16 

Nitrogen 0.80 0.29 

Chlorine 0 0 

Sulfur 0.22 0.15 

Oxygen 40.63 28.34 
24 

Pressurized biomass is then conveyed into the gasifier. Oxygen at 95% purity is produced from 

the air separation unit. A fixed 0.35 mass ratio of oxygen to biomass is used for the entrainedflow gasifier as reported by Henrich [18]. Steam addition to the gasifier is set at 0.48 mass ratio 

of steam to biomass in accordance with Probstein and Hicks [11] and explained further in 

Appendix C. This gasifier operates at a temperature of 1,300°C, meaning that equilibrium can be 

modeled according to Frey and Akunuri [3]. The reactions shown in equations 3–9 are modeled 

using equilibrium constants. 

ܥ ൅ 2ܪଶ ՞ ܪܥସ (eqn. 3)

2ܥ ൅ 1.5ܱଶ ՞ ܥܱ ൅ ܥܱଶ (eqn. 4)

ܥܱ ൅ ܪଶܱ ՞ ܥܱଶ ൅ ܪଶ (eqn. 5)

2ܥܱ ൅ ܱଶ ՞ 2ܥܱଶ (eqn. 6)

ܵ൅ܪଶ ՞ ܪଶܵ (eqn. 7)

0.5ܰଶ ൅ 1.5ܪଶ ՞ ܰܪଷ (eqn. 8)

(9 .eqn (ଶܪ ൅ ܱܵܥ ՞ ܵଶܪ ൅ ܱܥ

The LT scenario gasifier uses a 0.26 mass ratio of oxygen to biomass at a gasification 

temperature of 870°C. This ratio is developed from the data found in an IGT gasifier study by 

Bain [4]. In that study, Bain develops mass balances for an IGT gasifier operating with woody 

biomass. Steam addition to the gasifier is calculated using a 40/60 steam to oxygen mass ratio, 

consistent with experiments performed at Iowa State University using corn stover feedstock and 

a steam/oxygen-blown fluidized-bed gasifier. Low-temperature gasification cannot be modeled 

at equilibrium with or without approach temperatures for reactions. Instead, an elemental mass 

balance calculation and adjustment is performed to ensure all inlet and outlet streams are 

accounted for across the gasifier. For details on the LT gasifier mass balance calculation, see 

Appendix C. 

Yield from each gasifier is different. As Table 6 shows, hydrocarbons and tars are not produced 

in the high-temperature gasifier because of near equilibrium conditions. Also, more hydrogen 

formation occurs in the high-temperature gasifier as a result of the water-gas-shift reaction 

(equation 5) and because thermodynamically, nearly no methane, ethane, and ethylene are 

produced. The low-temperature gasifier, on the other hand, produces a significant amount of 

methane, ethane, and ethylene in the syngas, requiring downstream reforming. In the HT 

scenario, slag forms from the ash when the ash melts and flows on the inside walls. It is then 

collected at the bottom and removed for storage and subsequent waste removal. In accordance 

with Frey and Akunuri [3], it is assumed that 95% of the ash in the stover becomes slag while the 

rest becomes fly ash. 
25 

Table 6. Syngas Composition (Mole Basis) Leaving Gasifier for Gasification Scenarios Evaluated 

Component High Temperature 

(mole fraction) 

Low Temperature 

(mole fraction) 

Carbon Monoxide 0.264 0.240 

Hydrogen 0.310 0.200 

Carbon Dioxide 0.137 0.274 

Water 0.280 0.194 

Nitrogen 0.002 0 

Methane 6×10-6 0.055 

Ethane 0 0.0061 

Ethylene 0 0.013 

Ammonia 3.1×10-5 0.0094 

Hydrogen Sulfide 6.72×10-4 0.00112 

Carbonyl Sulfide 2.6×10-5 0 

Tar (Anthracene) 0 5.00×10-4

Oxygen 0 0 

Argon 0.006 0.006 

Directly after the low-temperature gasifier, initial syngas cleaning by cyclones captures char and 

ash. The cyclones are split into two trains because of high volumetric gas flow. Each train 

contains a medium efficiency cyclone, followed by high efficiency cyclones to capture 

particulates. Overall particulate removal efficiency for the cyclone area is 99%. Nearly 

particulate-free syngas travels to the more rigorous syngas cleaning in Area 300. Captured char 

in the LT scenario is collected and combusted in a fluidized-bed combustor providing energy for 

heating low-pressure steam, which is used for drying the stover. Syngas produced in the HT 

scenario contains fly ash, which is subsequently removed in a direct-water-quench unit. The 

combustion area in the HT scenario receives unconverted syngas from the fuel synthesis area, as 

char is not produced. For both scenarios, the combustor is assumed to operate adiabatically, 

resulting in an exit flue gas temperature of approximately 1,800°C. Hot flue gas heats 120°C 

steam to 200°C and loops to the stover drying area. 

Area 300 Syngas Cleaning 

After the initial particulate removal accomplished by the cyclones, the syngas still contains some 

particulates and all of the ammonia, hydrogen sulfide, and other contaminants. Area 300 removes 

these species using a cold-gas-cleaning approach, which is presently proven in many commercial 

configurations. Hydrogen sulfide and carbon dioxide, collectively known as acid gas, are 

absorbed via amine scrubbing. The LO-CAT hydrogen sulfide oxidation process is used to 

separate carbon dioxide from hydrogen sulfide with subsequent recovery of solid sulfur. In 

addition, the HT scenario contains a sour water-gas-shift (SWGS) process (sour because of the 

presence of sulfur), whereas the LT scenario situates the water-gas-shift (WGS) process directly 

upstream from the Fischer-Tropsch reactor. 

Because of a less than optimal hydrogen to carbon monoxide ratio from the gasifier, a WGS 

reaction is necessary at some point in the process to adjust to the optimum Fischer-Tropsch ratio 

of 2.1. Therefore, a significant WGS activity is required, so a sizable amount of carbon dioxide is 

produced. To keep that carbon dioxide from building up in downstream processes, the SWGS 
26 

reactor is located before the acid gas removal area. This SWGS unit operation is the most 

significant difference between the HT and LT scenarios in this area. 

In the HT scenario, the syngas arriving from the gasifier is cooled by direct-contact water quench 

to the operating temperature of the SWGS unit. In addition to cooling, the direct water quench 

removes all of the fly ash, sludge, and black water to prevent downstream plugging. At this 

point, a portion of the syngas is diverted to the SWGS unit, which is modeled at equilibrium 

conditions and has an exit gas temperature of 300°C. A ratio of 3:1 water to carbon monoxide is 

reached by addition of steam to the SWGS reactor. After the syngas is combined, the gas is 

further cooled to prepare for the acid gas removal. In the LT scenario, the direct quench unit 

condenses the syngas, removing approximately 90% of ammonia and 99% of solids. Tar is 

condensed in this unit and can be recycled back into the gasifier using a slurry pump, but this 

configuration is not modeled. A water treatment facility for the direct-quench effluent is not 

modeled but is accounted for in a balance-of-plant cost. 

The next step for cleanup is the removal of acid gas (carbon dioxide and hydrogen sulfide) 

through the use of an amine-based solvent in a chemical gas absorption system. At this point in 

the cleaning process, hydrogen sulfide and carbon dioxide content is approximately 900 ppm and 

30% on a molar basis, respectively. Sulfur must be removed to at least 0.2 ppm for FischerTropsch synthesis [31]. According to the Gas Processors Suppliers Association Engineering 

Data Book [44], amine-based systems are only capable of removing sulfur down to 4 ppm. 

Therefore, a zinc oxide guard bed is required to remove the difference. In this study, 20% 

concentrated monoethanolamine (MEA), capable of absorbing 0.4 mole acid gas per mole amine, 

is used as the absorbent. The process setup is based on a report by Nexant Inc. [27]. Hydrogen 

sulfide leaves the top of the absorber at 4 ppm and CO2 leaves at 2%, which is 99% and 90% 

removal, respectively. The clean syngas is now ready for polishing to final cleanliness 

requirements. A stripper is utilized to desorb the acid gas and regenerate the amine solution. 

Before the acid gas and amine solution enters the stripper, a heat exchanger raises the 

temperature to 90°C. 

Acid gas is brought to the LO-CAT sulfur recovery system to isolate hydrogen sulfide and 

convert it to solid sulfur. The LO-CAT system sold and owned by Gas Technology Products uses 

oxygen and a liquid solution of ferric iron to oxidize hydrogen sulfide to elemental solid sulfur 

[45]. This system is suitable for a range of 150 lb to 20 tons per day sulfur recovery and 100 ppm 

to 10% H2S concentration in sour gas, as reported by Nexant Inc. [27]. The sulfur production in 

this model is approximately 3 metric tpd with H2S concentration of approximately 150 ppm, 

which is within the reported ranges. First, the H2S is absorbed/oxidized, forming solid sulfur and 

water, while the ferric iron converts to ferrous iron. The second vessel oxidizes the ferrous iron 

back to ferric iron and the sulfur cake is removed while the iron solution is recycled back into the 

absorber [46]. The carbon dioxide gas stream from the absorber is split, and a portion is 

compressed and used in biomass pressurization while the rest is vented to the atmosphere. 
27 

Area 400 Fuel Synthesis 

Conversion from syngas to liquid fuel occurs in Area 400, the fuel synthesis area. The major 

operations in this area are zinc oxide/activated carbon gas polishing, steam methane reforming 

(SMR) (only in the LT scenario), water-gas-shift (only in the LT scenario), Fischer-Tropsch (FT) 

synthesis, hydrogen separation via pressure swing adsorption (PSA), FT products separation, and 

unconverted syngas distribution. Another major difference between the LT and HT scenarios is 

in this area. Area 400 in the LT scenario contains the water-gas-shift reaction and steam methane 

reformer, as recycle streams contain a high enough content of methane and ethylene to 

significantly accumulate and cause dilution. 

A compressor is the first operation in Area 400, boosting the pressure to 25 bar for FT synthesis. 

Then the syngas is heated to 200°C and passes through zinc oxide/activated carbon fixed-bed 

sorbent. This polishing guard bed acts as a barrier to any upstream non-normal contaminant 

concentrations and provides sulfur removal down to synthesis requirements. To limit 

downstream catalyst poisoning, the syngas stream must be cleaned of these components. 

Removal to 50 ppb sulfur is possible with zinc oxide sorbent [27]. To comply with reported 

requirements, the sorbent removes sulfur to approximately 200 ppb. In addition to sulfur, halides 

are removed by the sorbent. Syngas contaminant level requirements for Fischer-Tropsch 

synthesis are shown in Table 7. 

Table 7. Fischer-Tropsch Gas Cleanliness Requirements [31] 

Contaminant Tolerance Level 

Sulfur 0.2 ppm (200 ppb) 

Ammonia 10 ppm 

HCN 10 ppb 

Halides 10 ppb 

Methane, nitrogen, and carbon dioxide act as inert gases in the FT synthesis. At this point in the 

LT scenario, an SMR step is utilized. Syngas is heated to 870°C through a fired heater and 

passed through a reformer nickel-based catalyst to reduce methane, ethylene, and ethane content. 

It is assumed that the SMR can be modeled to operate at equilibrium. Steam is added to bring the 

steam to methane ratio to approximately 6.0, which at 870°C and 26 bar results in about 1.5% 

equilibrium methane content in the exit stream [47]. For the HT scenario, the SMR step is not 

necessary. The WGS reaction is now employed for the LT scenario to increase the H2:CO ratio. 

A portion of the gas is diverted through the fixed catalyst bed while the rest bypasses the reactor, 

similar to the SWGS unit in the HT scenario. 

The exiting H2:CO ratio after WGS is slightly above 2.1 in order for the excess hydrogen to be 

separated and used in the hydroprocessing area. A PSA process is employed to isolate a stream 

of hydrogen. Because only a small amount of hydrogen needs to be separated from the syngas 

stream for downstream use, a small percentage of the syngas is directed to the PSA unit. 

Hydrogen removal efficiency within the PSA unit is assumed to be 85%, and the unit produces 

pure hydrogen [41]. After the PSA, the syngas rejoins the main gas line and enters the FT 

reactor. 
28 

The Fischer-Tropsch synthesis reactor operates at 200°C and 25 bar using a cobalt catalyst 

according to equation 10. Per-pass carbon monoxide conversion in the reactor is set at 40%. The 

product distribution follows the Anderson-Schulz-Flory alpha distribution where chain growth 

factor, α, depends on partial pressures of H2 and CO and the temperature of the reactor, as 

reported by Song et al. [48] for cobalt catalyst and shown in equation 11, where ݕ is the molar 

fraction of carbon monoxide or hydrogen and ܶ݁݉݌ is the reactor operating temperature in 

kelvin. The reactor is based on a fixed-bed type reactor and that choice is reflected by the low 

per-pass CO conversion. 

ܥܱ ൅ 2.1ܪଶ ՜ െሺܪܥଶሻ െ ൅ܪଶܱ (eqn. 10)

ை஼ݕ · 0.2332൤ ൌ ߙ

ுଶݕ ൅ ை஼ݕ

൅ 0.6330൨ · ሾ1 െ 0.0039ሺܶ݁݉݌ െ 533ሻሿ (eqn. 11)

To ensure that the hydrocarbon product distribution leans toward the production of diesel fuel, 

the value of alpha should be at least 0.85 and preferably greater than 0.9, as shown in Figure 11. 

The reactor operating temperature needed to achieve a chain growth value of 0.9 is 

approximately 200°C. This produces 30 wt% wax in the FT products, requiring hydrocracking 

before addition to the final fuel blend. All exiting effluent is cooled to 35°C and the liquid water 

and hydrocarbons are separated in a gas/liquid knock-out separator. Unconverted syngas is split 

into four streams: direct recycle to the FT reactor, recycle to the acid gas removal area, purge to 

the combustor in Area 200, and to the gas turbine in the power generation area. The LT scenario 

does not include sending a syngas stream to the combustor in Area 200 because char is used. 

Overall CO conversion is 66% due to recycling syngas. The recycle ratio is approximately 1.95 

for both scenarios. 
29 

Figure 11. Fischer-Tropsch product distribution as a function of chain growth factor (ࢻ (using 

equation 11 [48] 

Area 500 Hydroprocessing 

FT products from the fuel synthesis area contain significant amounts of high-molecular-weight 

wax. Hydrogen is required to crack these high-molecular-weight parrafins to low-molecularweight hydrocarbons. A product distribution is specified in Table 8, as detailed in Shah et al. 

[49]. It is assumed that the hydroprocessing area contains a hydrocracker for converting the wax 

fraction and a distillation section for separating naphtha, diesel, and lighter-molecular-weight 

hydrocarbons. Also, hydrogen is assumed to be recycled within this area as needed. Methane and 

propane are separated and used to fuel the gas turbine in the power generation area. The 

hydroprocessing area is modeled as a “black box.” 

Table 8. Hydroprocessing Product Distribution [49] 

Component Mass Fraction 

Methane 0.0346 

LPG (propane) 0.0877 

Gasoline (octane) 0.2610 

Diesel (hexadecane) 0.6167 

Area 600 Power Generation 

A gas turbine and steam turbine provide the means to produce the power that is required 

throughout the plant and the excess power that is generated for export. Unconverted syngas from 

Fisher-Tropsch synthesis and fuel gas from hydroprocessing are combusted in a gas turbine, 

producing hot flue gas and shaft work. The flue gas exchanges heat with water in a heat-recovery 

steam generator to produce steam for the steam turbines, which subsequently produce more shaft 

work. Electric generators attached to both the gas turbine and the steam turbine produce 

electricity from the shaft work. 

0.00

0.20

0.40

0.60

0.80

1.00

0.00 0.20 0.40 0.60 0.80 1.00

Weight Fraction

Chain Growth Factor, α

Weight Fraction of Alkanes across Chain Growth Factor Range

C1 = methane, C2 = ethane, C3 = propane, etc.

C1-C4

C5-C11

C12-C19

C20-C120
30 

Area 700 Air Separation 

Because 95% purity oxygen is used for both scenarios, a cryogenic ASU is employed rather than 

purchasing oxygen. A two-column cryogenic oxygen/nitrogen separation system is employed 

with subsequent oxygen compression and nitrogen vent. Air pre-cooling is accomplished by 

exchanging heat with exiting nitrogen. This area requires a significant amount of power, as 

explained in the results section, which is provided by the power generation area. 

Methodology for Economic Analysis 

Capital investment and PV of each scenario are determined by finding all equipment costs and 

operating costs for the construction and operation of a plant for 20 years. Total capital 

investment is based on the total equipment cost plus additional installation costs and indirect 

costs (such as engineering, construction, and contingency costs). Annual operating costs are 

determined and a discounted cash flow rate of return analysis is developed. PV per unit volume 

of fuel is determined at a net present value of zero and a 10% internal rate of return. The major 

economic assumptions used in this analysis are listed in Table 9. A detailed list of assumptions 

can be found in Appendix A. 

Table 9. Main Economic Assumptions for nth Plant Scenarios 

Parameter Assumption 

Financing 100% equity 

Internal rate of return (after taxes) 10% 

General plant depreciation period 7 years (all areas except Area 600) 

Steam plant depreciation period 20 years (Area 600 only) 

Construction period 

2.5 years with total capital investment spent at 8%, 

60%, and 32% per year during years before 

operation 

Startup time 

0.5 years, where during that time revenues, variable 

operating costs, and fixed operating costs are 50%, 

75%, and 100% of normal, respectively. 

Income tax rate 39% 

Contingency 20% of fixed capital investment 

Electricity cost 5.4 cents/kWh 

Working capital 15% of fixed capital investment 

Land purchase 6% of total purchased equipment cost 

Plant availability 310 days per year (85%) 

Unit operations from the scenarios are sized and costs estimated using Aspen Icarus Process 

Evaluator software based on the Aspen Plus simulation data. Unique equipment costs for such 

equipment as the gasifier and Fischer-Tropsch synthesis reactor are estimated externally using 

literature references. Additionally, some equipment, such as the biomass dryer and lockhoppers, 

require literature references to determine the sizing, whereby their costs are subsequently 
31 

estimated using Aspen Icarus. The hydroprocessing plant area is modeled as a “black box” and 

therefore its costs are estimated as an overall scaled area cost from literature. 

The costs of each equipment item and process area are scaled based on a scaling stream and 

scaling size factor (݊) according to equation 12, where the size factor is between 0.6 and 1.0, 

depending on the equipment type. 

௪௡௘݁ݖܵ݅ ݁ܽ݉ݎݐܵ൤ כ ଴ݐݏ݋ܥ ൌ ௪௡௘ݐݏ݋ܥ

଴݁ݖܵ݅ ݁ܽ݉ݎݐܵ

൨

௡

(eqn. 12)

All purchased equipment costs determined via Aspen Icarus contain an installation factor that 

accounts for piping, electrical, and other costs required for installation. However, this installation 

factor tends to be significantly lower than metrics suggested by Peters et al. [50]. Therefore, 

rather than using the software-derived installation factors, an overall installation factor is applied 

to most equipment. A 3.02 overall installation factor is used, as suggested by Peters et al. for 

solid-liquid plants. Basically, the purchased equipment cost of a piece of equipment is multiplied 

by the installation cost to determine its installed cost. For the gasification unit, a 2.35 installation 

factor is used according to a National Energy Technology Laboratory study by Reed et al. [51]. It 

is assumed that all gas compressors receive a 1.2 installation factor, which is consistent with 

Aspen Icarus. The Chemical Engineering Plant Cost Index is used to bring the cost to 2007 U.S. 

dollars whenever a source for an estimated cost is from a previous year [52]. For multiple unit 

operations that operate in parallel or in trains, a train cost factor is applied. The reason for the 

factor, as reported by Larson et al. [38], is because those units share some piping, electrical, and 

other installation costs. It is applied as shown in equation 13, where ݊ is the number of units in 

the train and ݉ is the train factor with a value of 0.9. 

ݐݏ݋ܥ௧௥௔௜௡ ൌ ݐݏ݋ܥ௨௡௜௧ כ ݊) ௠eqn. 13)

Table 10 explains the methodology used to estimate capital investment. After total purchased 

equipment cost (TPEC) and total installed cost (TIC) are determined, indirect costs are applied. 

Indirect costs (IC) include engineering and supervision, construction expenses, and legal and 

contractors’ fees at 32%, 34%, and 23% of TPEC, respectively [50]. Project contingency is 

added as 20% of total direct and indirect cost (TDIC). TDIC is set as the sum of IC and total 

installed costs (TIC). Adding project contingency to TDIC gives the fixed capital investment. 

Total capital investment (TCI) is determined by adding working capital to fixed capital 

investment and thereby represents the overall investment required for each scenario. 
32 

Table 10. Methodology for Capital Cost Estimation for nth Plant Scenarios 

Parameter Method 

Total Purchased Equipment Cost (TPEC) Aspen Icarus Process Evaluator, 

references 

Total Installed Cost (TIC) TPEC Installation Factor 

Indirect Cost (IC) 89% of TPECa 

Total Direct and Indirect Costs (TDIC) TIC + IC 

Contingency 20% of TDIC 

Fixed Capital Investment (FCI) TDIC + Contingency 

Working Capital (WC) 15% of FCI 

Total Capital Investment FCI + WC 

a

 Indirect costs are broken down into engineering and supervision, construction expenses, and legal and 

contractors’ fees at 32%, 34%, and 23%, respectively, for a total of 89% of TPEC. 

Raw material costs are inflated to 2007 U.S. dollars using the Industrial Inorganic Chemical 

Index also used by Phillips et al. [39]. Annual variable operating costs are determined from 

material stream flows. Variable operating costs and their respective cost methods are shown in 

Table 11. Natural gas for use in the gas turbine to produce power during startup and backup 

periods is assumed to be employed 5% of the annual operating time. Solids disposal costs are for 

the handling and removal of ash in the LT scenario and slag in the HT scenario. Wastewater 

disposal cost is applied to the sludge and black water produced during direct syngas quench. 

Catalyst costs are not calculated on an annual basis because the catalysts for all reactors are 

assumed to be replaced every 3 years. Instead they are accounted for in the discounted cash flow 

analysis. 
33 

Table 11. Variable Operating Cost Parameters Adjusted to $2007 

Variable Operating Costs Cost information 

Feedstock $75/dry short ton 

LO-CAT Chemicals $176/metric ton of sulfur produced as reported 

in Phillips et al. [39] 

Amine Makeup $1.09/lb as reported in Phillips et al. and set as 

0.01% of the circulating rate [39] 

Process Steam $8.20/ton (Peters et al.) [50] 

Cooling Water $0.31/ton (Peters et al.) 

Hydroprocessing $4.00/barrel produced as reported by 

Robinson and Dolbear [53] 

Natural Gas (for backup) $6.40/thousand standard cubic feet as the 

average wellhead price for 2007 [54] 

Ash/Char Disposal $23.52/ton [39] 

Wastewater Disposal $3.30/hundred cubic feet [39] 

Electricity $0.054/kWha

Sulfur $40.00/ton [39] 

Fischer-Tropsch Catalyst 

(cobalt) 

$15/lb and 64 lb/ft3 density; applied on first 

operation year and then every three yearsa

Water-Gas-Shift Catalyst 

(copper-zinc) 

$8/lb and 900 kg/m3

; applied on first operation 

year and then every three years. Sour shift and 

normal WGS are assumed to operate with 

same catalysta

Steam Methane Reforming 

Catalyst (nickel-aluminum) 

$15/lb and 70 lb/ft3; applied on first operation 

year and then every three yearsa

Pressure Swing 

Adsorption $2/lba

a

 Assumed. 

Fixed operating costs include employee salaries, overhead and maintenance, and insurance and 

taxes. Salaries are calculated similarly to Phillips et al. [39], where employees include a plant 

manager, shift supervisors, lab technicians, maintenance technicians, shift operators, yard 

workers, and office clerks. The labor index developed by the Bureau of Labor Statistics [55] is 

used to adjust the labor cost to $2007. Overhead is calculated as 60% of total salaries; 

maintenance cost and taxes/insurance cost are both 2% of total installed equipment cost in 

accordance with Aden et al. [56]. 

For the DCFROR analysis, the capital investment is spent over a 2.5 year construction period, 

with 8% in the first half-year, followed by 60% and 32% for the second and third years. Working 

capital is applied in the year before operation and recovered at the end of the plant life. A 

standard modified accelerated cost recovery system (MACRS) is used, with the steam plant 

depreciating over 20 years and the rest of the plant over a 7-year period consistent with IRS 
34 

allowances. The project life is 20 years. Plant availability of 310 days per year (85%) is assumed 

and affects raw materials purchase as well as fuel production. The PV per gallon of gasoline 

equivalent is calculated for a set net present value of zero including a 10% internal rate of return. 

Methodology for Major Equipment Costs 

The software used for determining equipment costs is not capable of estimating every unit in this 

study. Some units such as the gasifiers and Fischer-Tropsch reactors are unique pieces of 

equipment that are underestimated if estimated as a simple vertical pressure vessel. Therefore, 

literature sources are used to help estimate sizes and costs of many units. The following section 

details a few of the more important units. 

The biomass dryer costs are estimated by determining the drying contact area. According to 

Couper [57], typical rotary dryers have a diameter of 6 feet and solids holdup of 8%. Assuming a 

bulk density of 100 kg/m3

 for ground stover and 1,000 kg/m3 for moisture in the stover, the 

resulting total surface area required for drying is 1,880 m2. The surface area provides enough 

information for estimating the costs because rotary dryer costs are estimated based on surface 

area in Aspen Icarus. Details on dryer sizing can be found in the Detailed Calculations section of 

Appendix C. 

The lockhopper system sizes are estimated by referring to a U.S. Department of Energy report 

completed by Combustion Engineering, Inc. [58], in which residence times and operating 

pressures are given. The biomass receiving bin, lockhopper, and feed bin costs are then estimated 

with Aspen Icarus. Details on lockhopper sizing can be found in the Detailed Calculations 

section of Appendix C. 

The high-temperature gasifier cost is estimated from Reed et al. [51]. The total bare erected cost 

(installed cost) of a train of eight high-temperature E-Gas gasifiers (2,500 metric tpd coal) 

including syngas cooling costs is $638 million ($2006). It is assumed that the syngas cooling 

accounts for 20% of that cost, and therefore the estimated installed cost in millions of $2006 for 

a 2,000 metric tpd high-temperature gasifier follows the formula in equation 14, resulting in $57 

million installed cost. 

ݐݏ݋ܥு்௚௔௦௜௙௜௘௥ ൌ ሺ638 · 80%ሻ

ܶܯ2000൤ · 8

൨ܶܯ2500

଴.଻

(eqn. 14)

Installed cost for a fluidized-bed gasifier is described in Larson et al. [38] and is calculated as 

shown in equation 15, where ݐݏ݋ܥ௔௚_଴௦௜௙௜௘௥ is $6.41 million ($2003), ܵݎݐ݁ܽ݉ ܵ݅ݖ݁ ଴is 41.7 

metric ton per hour, and ݊ is 0.7. The gasifier is evaluated at 300 short tpd because that appears 

to be the highest proven capacity for a GTI gasifier. Therefore, seven fluidized-bed gasifiers are 

used in parallel. It is assumed that the gasifier train follows the train cost formula (equation 13), 

resulting in $19 million installed cost. 

݁ݖܵ݅ ݁ܽ݉ݎݐܵ ൤ כ ௥௘௜௙௜௦଴_௚௔ݐݏ݋ܥ ൌ௜௙௜௘௥ ௦௚௔்௅ݐݏ݋ܥ

଴݁ݖܵ݅ ݁ܽ݉ݎݐܵ

൨

௡

(eqn. 15)
35 

In a similar manner, the FT reactor is estimated as described in Larson et al. [38], where base 

installed cost is $10.5 million ($2003), base sizing value is 2.52 million scf/h of synthesis gas 

flow, and sizing exponent is 0.72. An installation factor of 3.6 is assumed for the FT reactor as 

found in Peters et al. [50] for liquid production plants. This allows the purchased cost of the unit 

to be back-calculated. 

The acid gas removal (AGR) area cost is evaluated using information from Phillips et al. [39], 

following equation 12 where the base stream size is 4,000 short tpd and the base cost is $5.45 

million. The stream size is the mass flow of the synthesis gas entering the AGR as the sum of 

fresh syngas from gas scrubbing and unconverted syngas from the fuel synthesis area. 

Capital investment for the hydroprocessing area is found in Robinson et al. [53]. That study 

reports a volumetric unit cost of $4,000 per barrel per standard day. Assuming the typical 

hydroprocessing refinery produces 25,000 barrels per day, the base cost, C0, is $100 million. 

Assuming a scaling exponent of 0.65, the cost of Area 500 is found using equation 12. The cost 

details for both of the gasifiers, the AGR area, the FT reactor, and the hydroprocessing area can 

be found in the Detailed Calculations section of Appendix C. 

Methodology for Sensitivity Analysis 

Sensitivity parameters are chosen to reflect the change in PV. The parameters are either 

economic or process parameters. The sensitivity bounds are chosen based on what is expected to 

be observed in the construction and operation of a biomass-to-liquids production plant. The 

chosen favorable, baseline, and unfavorable sensitivity variables are shown in Table 12. 

Table 12. Sensitivity Parameters for nth Plant Scenarios 

Parameter Favorable Baseline Unfavorable 

Availability (hours/year) 8,000 7,446 7,000 

Balance of Plant (% of TPEC)a 8 12 16 

Catalyst Cost (%)b 50 100 200 

Catalyst Lifetime (years) 5 3 1 

CO Conversion in FT Reactor (%) 30 40 50 

Compressor Install Factor 1.0 1.2 3.0 

Contingency (% of TDIC)c 10 20 30 

Feedstock Cost ($/dry short ton) 50 75 100 

Feedstock Moisture (% wet) 20 25 30 

Price of Electricity (¢/kWh) 7.0 5.4 3.0 

Total Capital Investment (% of 

baseline) 70 100 130 

a

 TPEC=total purchased equipment cost. 

b

 All catalyst costs are varied over this range. 

c

 TDIC=total direct and indirect cost. 
36 

Methodology for Pioneer Plant Analysis 

Economic analysis is based on an nth plant design, and before a project is undertaken the pioneer 

(1st) plant cost is important to estimate. We use a method begun by the RAND Corporation that 

estimates pioneer plant costs and plant performance. Using this methodology, two main areas of 

the nth plant economic analysis are adjusted: capital investment and plant performance. Through 

a series of parameters, a cost inflation factor is generated to inflate the capital investment. In 

addition, a plant performance factor is calculated that reduces the fuel sales, feedstock purchase, 

and variable operating costs for the first several years that the plant is in operation. Each year, 

the plant performance factor is increased until full performance is attained. For the purpose of 

determining a range of pioneer plant costs, baseline, optimistic, and pessimistic values are 

chosen. The details of the RAND methodology can be found in Merrow et al. [59]. The 

following section explains the reasoning behind the parameters chosen for the scenarios. 

Cost growth and plant performance factors are calculated as shown in equations 16 and 17 in 

accordance with Merrow et al. [59]. 

ܵܧܫܶܫܷܴܲܯܫ כ 0.02125 െ ܹܧܶܰܥܲ כ 0.00297 െ 1.1219 ൌ ݄ݐݓ݋ݎܩ ݐݏ݋ܥ

ܵܵܧܰܧܸܫܷܵܮܥܰܫ כ 0.00111 ൅ ܻܶܫܺܧܮܲܯܱܥ כ 0.01137 െ

ܱܰܫܶܫܰܫܨܧܦ ܶܥܧܬܴܱܲ כ 0.06351 െ

(eqn. 16)

4.12 െ ܳܵܧܮܣܤ כ 0.33 ൅ ܲܵܧܹܵܶܧܰ כ 9.69 െ 85.77 ൌ ݂.ݎܲ݁ ݐ݈ܲܽ݊

(17 .eqn (ܵܦܫܮܱܵ כ 17.91 െ ܧܵܶܣܹ כ

The variables used in determining pioneer plant performance (equation 17) are defined as 

follows: 

NEWSTEPS (0+): The feedstock handling area is chosen as a new step because of the 

large scale, which has not been demonstrated with biomass. The gasifier and solids 

feeding are also included as a new step because a pressurized biomass feeding system has 

not been demonstrated at a commercial scale except for in limited campaigns. 

BALEQS (0 to 100): The mass and energy balances cannot be validated with current 

plant data, so a value of zero is chosen. 

WASTE (0 to 5): Waste streams for gasification include scrubber sludge, black water, 

gasifier slag, fly ash, and sulfur. The scrubber sludge and black water require chemical 

treatment and the sulfur requires special handling. A mid-range value of 2.5 is chosen. 

SOLIDS (0 or 1): Solids are present; therefore a value of 1 is used. 

Variables used in determining pioneer plant cost growth (equation 16) are defined as follows: 

PCTNEW (0 to 100%): The percentage cost of the gasifier, solids pressurizing, and solids 

feeding out of the total purchased equipment cost. 
37 

IMPURITIES (0 to 5): There are two major recycle streams in the gasification process, 

and there is the possibility of inert component buildup. There is also a potential for 

equipment corrosion due to sulfur components, hydrogen chloride, and hydrogen, so a 

value of 4 is assigned. 

COMPLEXITY (0+): There are nine continuously linked steps in the gasification 

process. These include feedstock handling, solids feeding, gasification, amine scrubbing, 

sour water-gas-shift, pressure swing adsorption, Fischer-Tropsch synthesis, 

hydroprocessing, and air separation. 

INCLUSIVENESS (0 to 100): Land costs and startup costs are considered in the TCI; 

however, they have not been rigorously investigated. A value of 33% is used. 

PROJECT DEFINITION (2 to 8): The gasification platform is considered to be in the 

study design stage so a value of 7 is assigned. 

The factors are applied to the capital investment and plant performance as shown in equations 18 

and 19. Expenses and revenues affected by the plant performance factor are fuel sales, feedstock 

purchase, co-product credits, and variable operating costs. 

௛௧௡ܫܥܶ ൌ௣௜௢௡௘௘௥ ܫܥܶ

(18 .eqn (݄ݐݓ݋ݎܩ ݐݏ݋ܥ

The ݐݏ݋ܥ ݐݓ݋ݎܩ݄ factor causes the TCI of the pioneer plant to increase from the nth plant cost. 

ݐݏ݋ܥ௥௘௘௡௢௜௣ሺݐሻ ൌ ݐݏ݋ܥ௡௧௛ሺݐሻ כ ሺ݈ܲܽ݊ݐ ܲ݁ݎ݂. ൅ 20 כ ሺݐെ1ሻሻ

100 (eqn. 19)

ݐݏ݋ܥ௡௧௛ሺݐሻ is the nth plant expense or revenue at year ݐ. The plant performance factor is applied 

at year 1 and increases by 20% each year until 100% performance is reached. The chosen 

parameters and calculated factors for baseline, optimistic, and pessimistic scenarios are shown in 

Table 13. 
38 

Table 13. Pioneer Plant Analysis Parameters and Factors 

Parameter Baseline Optimistic Pessimistic Range 

NEWSTEPS 2 1 3 0+ 

BALEQS 0 0 0 0–100 

WASTE 4 3 5 0–5 

SOLIDS 1 1 1 0 or 1 

Plant Performance 38.18 49.93 22.31 0–100 

 

PCTNEW 19 (9)a 10 (5)a 25 (20)a 0–100 

IMPURITIES 4 3 5 0–5 

COMPLEX 9 6 12 0+ 

INCLUSIV. 33 50 0 0–100 

PROJ. DEF. 7 6 8 2–8 

Cost Growth (HT) 0.47 0.63 0.30 0–1 

Cost Growth (LT) 0.50 0.65 0.31 0–1 

a

 Value in parentheses is value chosen for LT scenario. 
39 

Results and Discussion 

Process Results 

With a higher carbon conversion efficiency (see mass and energy balance discussions below), the 

HT scenario produces 61.0 GGE of fuels per dry metric ton of feedstock compared to 47.2 

GGE/metric ton for the LT scenario. Resulting annual fuel outputs are 41.7 million GGE/yr for 

the HT scenario and 32.3 million GGE/yr for the LT scenario. 

Along with lower fuel yield, the LT scenario consumes less power (Table 14). The LT scenario 

and HT scenario total power usage are 15 and 22 MW, respectively. Major contributions to this 

result are a lower grinder power due to less strict biomass size requirement, lower pressurized 

oxygen consumption in the gasifier, and generally lower downstream mass flow rates throughout 

the plant for the LT scenario. Lower syngas yield also means that there is less unconverted 

syngas and fuel gas from the hydroprocessing area available for the gas turbine. Therefore, the 

LT scenario generates 31 MW compared to 36 MW generated by the HT scenario. Due to 

unoptimized flow rates of the recycle streams, the LT scenario actually generates a net 16 MW 

of power, which is more than the 14 MW produced in the HT scenario. Reducing the net power 

generation is achievable by increasing the recycle ratio and thereby increasing conversion, but a 

consequence is higher flow rates and therefore larger and more expensive equipment. The focus 

of this study is to produce liquid fuels. However, procedures to optimize recycle ratios, 

equipment sizes, and fuel production rates are not within the scope of this study and are not 

undertaken. 

Table 14. Power Generation and Usage 

Power (MW) HT Scenario LT Scenario 

USAGE 

Chopper 0.50 0.50 

Grinder 2.96 1.10 

Lockhopper System 0.18 0.18 

Lean Amine Solution Pump 0.86 0.69 

Syngas Booster 

Compressor 1.25 0.96 

PSA Compressor 0.15 0.11 

Recycle Compressor 0.39 0.29 

Hydroprocessing Area 2.24 1.73 

Oxygen Compressor (ASU) 3.61 2.80 

Air Compressor (ASU) 7.94 6.31 

Sour Gas Shift Steam 

Compressor 1.59 0 

CO2 Compressor 0.39 0.39 

Total Usage 22.06 15.06 

GENERATION 

Gas Turbine 26.25 21.02 

Steam Turbine 9.63 10.40 

Total Generated 35.88 31.42 

Net Export 13.82 16.36 
40 

An energy balance of the scenarios shows that the biomass-to-fuels efficiency for the LT and HT 

scenarios is 39% and 50% on a lower-heating-value (LHV) basis, respectively (Table 15). When 

the net electricity is added, the efficiencies are 43% and 53% on an LHV basis, respectively. The 

LT scenario is expected to be lower because mass and energy loss occurs in the production and 

removal of char and tar. Char and tar energy loss sums to 7.5% of the energy in the biomass. In 

this scenario, char is combusted in a fluidized-bed combustor to provide heat for biomass drying. 

Biomass drying in the HT scenario is accomplished by a syngas purge. However, the most 

significant energy loss in the LT scenario, about 25%, occurs across the gasifier. One reason for 

higher energy loss in the LT scenario is because conversion efficiency increases with increasing 

operating temperature. A second reason is the loss of energy during the cooling of the syngas 

after the gasifier. More effective capture of the energy in the hot syngas would increase the 

overall energy efficiency. 

High exothermicity of the FT reaction causes a significant portion of the chemical energy in the 

syngas to leave as thermal energy in both scenarios. A higher loss across the FT reactor is 

observed in the HT scenario due to higher flow rates. Energy closure, as shown in Table 15, is 

approximately 90% for both scenarios. It is assumed that the last 10% is due mostly to heat loss 

from the cooling of the syngas by direct quench rather than capturing the heat and raising steam. 

Table 15. Overall Energy Balance on LHV Basis 

HT Scenario LT Scenario 

IN

Biomass 100.0 100.0 

OUT

Fuel 49.7 38.5 

Net Electricity 3.5 4.2 

Power Gen Losses 4.2 3.1 

FT Reactor Losses 16.2 12.5 

Gasifier Losses 12.1 24.9 

Char - 6.3 

Tar - 1.2 

Syngas Purge 1.8 - 

Totala 87.5 90.7 a The balance of energy is assumed to come from various heating and cooling losses. 

A carbon balance analysis shows that 26% and 34% of the carbon in the biomass is passed on to 

the fuels for the LT and HT scenarios, respectively (Table 16). Approximately 99% of the carbon 

is accounted for. Major carbon losses include carbon dioxide flue gases, LO-CAT venting, and 

lockhopper venting. Char leaving the LT scenario is accounted for in the A200 flue gas because 

the char is combusted for process heat. Also, because the LT scenario produces low-molecularweight hydrocarbons in the gasification process, a small fraction is dissolved in the liquid 

effluent of the wet scrubber. Carbon dioxide also dissolves in the wet scrubber effluent stream. 

Another carbon loss comes from the hydrocarbons that dissolve in the acid gas removal area. 
41 

Table 16. Overall Carbon Balance 

HT Scenario LT Scenario 

 kmol/h % kmol/h % 

IN

Biomass 3,281 100 3,281 100 

OUT

Fuel 1,111 34 862 26 

A300 CO2 Vent 1,458 45 1,294 39 

A600 Flue Gas 334 10 302 9 

A200 Flue Gas 39 1 227 7 

Lockhopper Vent 159 5 162 5 

Wet Scrubber 

Effluent 154 5 318 10 

Tar 0 - 35 1 

Dissolved 

Hydrocarbons 0 - 46 1 

Total 3,257 99 3,245 99 

Throughout both scenarios, steam and cooling water are required as utilities. Because a pinch 

analysis (a method to optimize heat exchange) is not undertaken for this study, integration of the 

heat streams is not optimized. Therefore, it is assumed that the resulting heating and cooling 

requirements within the model represent steam and cooling water utilities, whereby they are 

recycled at a ratio of 9:1. In other words, fresh steam and cooling water utility input to the 

scenarios are calculated at 10% of the required circulating rate. 

Cost Estimating Results 

Capital and Operating Costs for nth Plant 

The breakdown of costs by area and resulting total capital investment is shown in Table 17. Total 

capital investment amounts for the HT and LT scenarios are $606 million and $498 million, 

respectively. Major areas of investment are the gasification area in the HT scenario and the fuel 

synthesis area in the LT scenario. Moreover, these two areas contain significant differences in 

capital investment between the scenarios. The installed cost of the entrained-flow gasifier is 

significantly higher than the cost of the fluidized-bed gasifiers, even when seven of the latter are 

used in parallel. Area 400 costs for the LT scenario are higher than for the HT scenario due to the 

steam methane reformer and additional heat-exchange equipment required for the high 

operational temperature. A significant portion of the capital cost is due to gas compression, such 

as the air compressor in the air separation unit and the syngas booster compressor. Due to high 

purchase costs, compressors make up approximately 18% of the TPEC for each scenario. 

Detailed accounting of equipment found in each process area can be found in Appendix B. 
42 

Table 17. Capital Investment Breakdown for nth Plant Scenarios 

HT Scenario LT Scenario 

Area 

Installed Cost Installed Cost 

$MM % $MM % 

A100: Preprocessing 23 7 23 9 

A200: Gasification 68 22 28 11 

A300: Syngas Cleaning 34 11 29 12 

A400: Fuel Synthesis 49 16 59 23 

A500: Hydroprocessing 33 11 30 12 

A600: Power Generation 46 15 39 15 

A700: Air Separation Unit 24 8 20 8 

Balance of Plant 33 11 27 11 

Total Installed Cost 309 254 

Indirect Cost 130 107 

Total Direct and Indirect 

Cost 439 361 

Contingency 88 72 

Fixed Capital Investment 527 433 

Working Capital 79 65 

Total Capital Investment 606 498 

Annualized costs for operation of the plant are shown in Table 18. The percentage displayed also 

represents percentage of PV. The largest annual incurred costs for both scenarios are the average 

return on investment and feedstock purchase. Utilities such as steam and cooling water are 

higher for the LT scenario due to heating and cooling of the syngas before and after the SMR and 

steam input to the SMR. Waste disposal costs are equal because equal amounts of ash or slag are 

by-products of both plants. Annual hydroprocessing area costs and income taxes are higher for 

the HT scenario because of its higher fuel production rate. Fixed costs and capital depreciation 

are also higher due to higher TCI. 

Catalyst costs are not determined on an annual basis because catalysts are assumed to be 

replaced only every 3 years. Table 19 contains catalyst replacement costs. The catalyst cost for 

the ZnO guard bed and PSA unit are equal across the scenarios because the volumes of the units 

are assumed to be the same. The FT catalyst for the HT scenario is significantly more expensive 

because of a higher gas flow rate and hence, higher catalyst use. Using a DCFROR analysis, the 

PVs at a net present value of zero for the LT and HT scenarios are $4.83/GGE and $4.27/GGE, 

respectively. Further detail of the yearly cash flow of the life of the plant can be found in 

Appendix B. 
43 

Table 18. Annual Operating Cost Breakdown for nth Plant Scenarios 

HT Scenario LT Scenario 

$MM % $MM % 

Average Return on 

Investment 58.2 32.7 48.3 31.0 

Feedstock 51.3 28.9 51.3 32.9 

Capital Depreciation 26.3 14.8 21.7 13.9 

Average Income Tax 21.9 12.3 18.0 11.6 

Fixed Costs 14.3 8.1 12.4 8.0 

Hydroprocessing 4.4 2.5 3.0 2.0 

Steam 2.7 1.5 3.5 2.2 

Cooling Water 2.3 1.3 3.5 1.6 

Waste Disposal 1.5 0.3 1.5 0.3 

Other Raw Material Costs 1.4 0.8 1.3 0.8 

Co-product Credits (5.6) -3.1 (6.6) -4.2 

Table 19. Catalyst Replacement Costs for Both Scenarios (3-Year Replacement Period) 

Catalyst HT 

scenario 

LT 

scenario 

Water-Gas-Shift 

(copper-zinc) $115,000 $105,000 

Steam Reforming 

(nickel-aluminum) N/A $103,000 

ZnO Guard Bed $424,000 $424,000 

PSA Packing $497,000 $497,000

Fischer-Tropsch 

(cobalt) $7,687,000 $6,128,000 

Sensitivity Results for nth Plant 

The results of sensitivity analysis are summarized in Figure 12 and Figure 13 for the HT and LT 

scenarios, respectively. Total capital investment and feedstock purchase cost have the highest 

effect on the PV at approximately ±$0.80 and ±$0.40 per GGE, respectively, for both scenarios. 

Due to the high percentage of equipment cost for compressors, the compressor installation factor 

has a very high effect on PV as well. When the compressor installation factor is increased to 3.0, 

which is the usual installation factor for most of the equipment, the PV increases by $0.71 and 

$0.78 per GGE for the LT and HT scenarios, respectively. Parameters with a lesser, but still 

significant, effect are the contingency factor (as percentage of total direct and indirect costs) and 

plant availability, both with an effect of approximately ±$0.20 per GGE. Parameters with the 

least effect are generally characteristic of the process rather than of the economics. For example, 

catalyst life, feedstock moisture, and carbon monoxide conversion in the FT reactor affect the PV 

less than ±$0.15 per GGE. 
44 

Figure 12. Sensitivity results for HT nth plant scenario 

Figure 13. Sensitivity results for LT nth plant scenario 

Additionally, the size of the plants can be varied by feedstock input rate. The effects of plant size 

on PV and TCI are shown in Figure 14 and Figure 15, respectively. When the plant size is 

reduced to 500 metric ton per day, the two scenarios approach equal PV. Also, as the plant size is 

reduced from the baseline, the difference in capital investment decreases. As the plant size 

increases past the baseline, the slope of the PV levels out, suggesting that the benefits of lower 

PV may not be worth the significant increase in capital cost (Figure 14). 

4.23

4.23

4.23

4.23

4.24

4.18

4.08

4.07

4.19

3.85

3.47

4.32

4.33

4.34

4.35

4.40

4.36

4.45

4.46

4.98

4.69

5.05

$3.25 $3.50 $3.75 $4.00 $4.25 $4.50 $4.75 $5.00 $5.25

Per pass CO Conversion in FT 50:40:30 (%)

Electricity Price 7.0:5.4:3.0 (¢/kWh)

Feedstock Moisture 20:25:30 (%wet)

Catalyst Cost 50:100:200 (%)

Catalyst Life 5:3:1 (year)

BOP 8:12:16 (% of TPEC)

Availability 8000:7446:7000 (hr/yr)

Contingency 10:20:30 (% of TDIC)

Compressor Install Factor 1.1:1.2:3.0

Feedstock Cost 50:75:100 ($/ton)

Total Capital 70:100:130 (%)

Product Value (2007$/GGE)

4.79

4.79

4.78

4.77

4.80

4.73

4.63

4.62

4.74

4.29

3.99

4.88

4.91

4.90

4.92

4.98

4.93

5.02

5.04

5.61

5.37

5.68

$3.75 $4.00 $4.25 $4.50 $4.75 $5.00 $5.25 $5.50 $5.75

Per pass CO Conversion in FT 50:40:30 (%)

Catalyst Cost 50:100:200 (%)

Feedstock Moisture 20:25:30 (%wet)

Electricity Price 7.0:5.4:3.0 (¢/kWh)

Catalyst Life 5:3:1 (year)

BOP 8:12:16 (% of TPEC)

Availability 8000:7446:7000 (hr/yr)

Contingency 10:20:30 (% of TDIC)

Compressor Install Factor 1.0:1.2:3.0

Feedstock Cost 50:75:100 ($/ton)

Total Capital 70:100:130 (%)

Product Value (2007$/GGE)
45 

Figure 14. The effect of plant size on product value (per gallon of gasoline equivalent) for nth plant 

scenarios 

Figure 15. The effect of plant size on total capital investment for nth plant scenarios 

Pioneer Plant Analysis Results 

The total capital investment for a base-case pioneer plant is expected to double from the nth plant 

scenario, as detailed in Table 20. PV for a base case pioneer plant is estimated to increase to 

$7.60 and $8.10 per GGE for the LT and HT scenarios, respectively. Table 20 also shows 

estimates for the optimistic and pessimistic cases. An important observation is that the pioneerplant PV for the LT scenario is actually lower than for the HT scenario. The reason behind this 

inverted result is because of the higher capital cost inflation (cost growth factor) in the HT 

scenario due to higher gasification area capital costs. 

0.00

1.00

2.00

3.00

4.00

5.00

6.00

7.00

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000 5500 

Product value ($/GGE)

Plant Size (metric ton/day)

HT Scenario

LT Scenario

0

200

400

600

800

1000

1200

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000 5500 

Total capital investment ($MM)

Plant Size (metric ton/day)

HT Scenario

LT Scenario

Base 

Base 
46 

Table 20. Pioneer Plant Analysis Results 

Analysis 

HT Scenario LT Scenario 

TCI ($MM) PV ($/GGE) TCI ($MM) PV ($/GGE) 

nth Plant 606 4.27 498 4.83 

1st Plant Base 1,400 8.10 1,080 7.60 

1st Plant 

Optimistic 1,030 6.30 830 6.30 

1st Plant 

Pessimistic 2,200 12.60 1,740 11.50 

Comparison with Previous Techno-Economic Studies 

Two previous BTL studies that specifically use biomass feedstock, low-temperature gasification, 

and Fischer-Tropsch synthesis technology are Tijmensen et al. [40] and Larson et al. [38]. In 

order to compare with these studies, major economic and process parameters from the present nth

plant LT scenario are adjusted to reflect similar values of the previous studies. First, the plant 

size of the present study is adjusted to increase equipment costs and raw materials purchases. As 

a result, the annual biomass input and TCI are affected. Second, availability in hours per year, 

rate of return, cost year, and feedstock cost are adjusted. The combined effect of all adjusted 

parameters causes the present study’s product value to reflect the comparison study. 

A comparison to the IGT-R scenario (which employs a low-temperature IGT gasifier and a steam 

methane reformer) in Tijmensen et al. shows that fuel product value is higher in the present 

study, as summarized in Table 21. Of all the scenarios developed by Tijmensen et al., the IGT-R 

scenario is most similar to the present study because of the reformer. The IGT-R scenario has a 

TCI of $387 million, a feedstock cost of $33 per short ton, and a product value of $1.90/GGE. 

An important characteristic of the Tijmensen et al. study is that it does not include a 

hydroprocessing area. Therefore, it might be expected that the TCI would be higher for the 

present study, which includes hydroprocessing. However, that is not the case, as the TCI of the 

present study using Tijmensen et al. parameters is $339 million, which is lower than the reported 

$387 million of the earlier study. Another important observation is that the annual fuel 

production for the present study, with adjusted parameters, is 30.2 million gallons per year, 

compared to 39.8 million gallons per year of FT products reported by Tijmensen et al. One 

reason for lower annual fuel production in the present study is because of a loss during 

hydroprocessing. Therefore, due to lower annual fuel production and hence lower fuel revenue, 

the present study has a higher product value compared to Tijmensen et al. 
47 

Table 21. Comparison of nth Plant LT Scenario to Tijmensen et al. Study IGT-R Scenario [40] 

Parameter 

Tijmensen et al. 

Study (IGT-R 

Scenario) 

Present Study nth

Plant LT 

Scenario 

Present Study w/ 

Tijmensen et al. 

Parameters 

Plant Size (dry tons/day) 1,920 2,205 1,920

Annual Biomass Input (tons) 640,000 684,100 640,000 

Total Capital Investment ($MM) 387 498 339

Availability (hour/year) 8,000 7,446 8,000

Rate of Return (%) 10 10 10

Cost Year 2000 2007 2000

Feedstock Cost ($/short ton) 33.00 75.00 33.00

Efficiency (%, LHV, incl. elec.) 50.1 42.7 42.7

Fuel Yield (MMGGE/yr) 39.8 32.3 30.2

Product Value ($/GJ) 16.50 39.80 25.17

Product Value ($/GGE) 2.00 4.83 3.05

A comparison to the FT-OT-VENT scenario (which is low-temperature gasification with carbon 

dioxide vent and once through FT synthesis) reported by Larson et al. is summarized in Table 22. 

In a similar fashion to the previous comparison, the parameters are adjusted to approximate the 

comparison study. Some important observations are made from this comparison. First, the TCI of 

the present study with adjusted parameters is significantly higher. Second, the net electricity is 

significantly lower for the present study. Third, the PV is significantly higher for the present 

study. Essentially, the Larson et al. study generates more revenue from selling electricity and 

recovers the capital investment in less time. In addition, annual operating costs for the Larson et 

al. study are lower than for the present study. Therefore, because of lower revenue from 

electricity and higher annual operating costs, the present study has a higher fuel product value 

when compared on a similar basis to Larson et al. 

 
48 

Table 22. Comparison of nth Plant LT Scenario to Larson et al. Study FT-OT-VENT Scenario [38] 

Parameter 

Larson et al. 

Study (FT-OTVENT Scenario) 

Present Study nth

Plant LT 

Scenario 

Present Study w/ 

Larson et al. 

Parameters 

Plant Size (dry tons/day) 5,000 2,205 5,001

Annual Biomass Input (tons) 1,458,000 684,000 1,459,000 

Total Capital Investment 

($MM) 541 498 678

Availability (hour/year) 7,000 7,446 7,000

Debt/Equity (% Equity) 60 100 60

Rate of Return (%) 12 10 12

Cost Year 2003 2007 2003

Electricity Price (cents/kWh) 4.0 5.4 4.0

Net Electricity (MW) 207 16.3 37.1

Feedstock Cost ($/short ton) 46.00 75.00 46.00

Plant Yield (MMGGE/yr) 63.3 32.3 68.9

Product Value ($/GJ) 15.25 39.80 26.80

Product Value ($/GGE) 1.85 4.83 3.25

Summary of nth Plant Scenarios 

The HT scenario requires more power and capital investment, yields more fuel per ton of 

feedstock, and subsequently produces more fuel per year compared to the LT scenario. The total 

capital investment requirements for the LT and HT scenarios are $498 million and $606 million, 

respectively. Despite higher capital investment for the HT scenario, the product value (PV) is 

lower. PVs for the LT and HT scenarios are $4.83 and $4.27 per gallon of gasoline equivalent, 

respectively. The main reason for a lower PV is because of increased fuel revenue. The main nth

plant scenario results are shown in Table 23. A detailed summary of costs can be found in 

Appendix B. 

Table 23. Main Scenario nth Plant Results 

Scenario TCI 

($MM) 

TPEC 

($MM) 

Fuel Yield 

(GGE/metric 

ton) 

Annual Fuel 

Output 

(MMGGE/yr) 

Net 

Electricity 

Export 

(MW) 

PV 

($/GGE) 

High 

Temperature 605.9 145.7 61.0 41.7 13.8 4.27 

Low Temperature 498.3 120.4 47.2 32.3 16.4 4.83 

. 
49 

Conclusions 

This analysis compares capital and operating cost for two biomass-to-liquids scenarios: hightemperature (HT) gasification and low-temperature (LT) gasification. The selection of these 

scenarios allows for direct comparison between two modes of gasification: slagging and nonslagging. The slagging, entrained-flow gasifier employed for the HT scenario results in higher 

plant costs (about 20%) than the LT scenario, which employs a fluidized bed gasifier. On the 

other hand, the higher carbon conversions for the HT gasifier result in a lower PV compared to 

the LT scenario. A biomass-to-liquids plant is expected to produce fuels costing in the range of 

$4–$5 per gallon gasoline equivalent using present gasification and Fischer-Tropsch synthesis 

technology. The factors chiefly responsible for this relatively high PV are feedstock costs and 

investment return on the capital to build a $500 million to $650 million plant to process 2,000 

metric tons of biomass per day. A pioneer plant analysis estimates that the total capital 

investment for a pioneer plant would double and PV would increase by approximately 60% from 

the nth plant scale. This uncertainty suggests that economics are still a major challenge for 

biomass-to-liquids production plants. 

Parameters with the most sensitive effect on PV are total capital cost, feedstock purchase cost, 

and compressor installation factor, affecting the PV ±$0.40–$0.80 per gallon. Less expensive 

biomass feedstock that is lower in ash content than that used in the present study will have higher 

fuel yield and have the potential to significantly decrease PV. Gas compression is a major 

portion of capital investment, and sensitivity analysis shows that installation costs of 

compressors have a high effect on PV. Factors with little effect on the PV are mostly related to 

the process, such as carbon monoxide conversion in the FT reactor, feedstock inlet moisture, and 

catalyst lifetime. 

Because of time and resource constraints, the techno-economic study presented includes a few 

shortcomings. The process configuration is not fully optimized for heat integration. While some 

recycle streams are included, a complete heat exchange network for heat recovery is not 

conceptualized. In addition, some areas such as FT product separation and hydroprocessing are 

not modeled rigorously and can be improved with detailed mass and energy flows. Further 

studies can benefit by optimizing heat exchange networks and modeling in detail areas such as 

FT product separation and hydroprocessing. More detailed modeling will likely result in changes 

to economic components such as equipment cost, installation factors, and contingency. However, 

because increasing model detail may result in tradeoffs between the different economic 

components, the net effect of increased model detail on PV is difficult to predict. 

This analysis tracks PV based on commercial technology for which sufficient public domain data 

existed in 2007. In order to evaluate the economics of biofuels produced from gasification based 

on future scenarios, particular attention will be needed on the most sensitive parameters—

feedstock cost and capital cost. Attention should also focus on commercial biomass-to-fuels 

gasification plants that come online. As these plants do come online, cost growth and plant 

performance factors will improve, thereby decreasing the pioneer plant PV. 
50 

References 

1. Kazi, F.K.; Fortman, J.A.; Anex, R.P.; Kothandaraman, G.; Hsu, D.D.; Aden, A.; Dutta, 

A. Techno-Economic Analysis of Biochemical Scenarios for Production of Cellulosic 

Ethanol. NREL/TP-6A2-46588. Golden, CO: National Renewable Energy Laboratory, 

2010. 

2. Wright, M.M.; Satrio, J.A.; Brown, R.C.; Daugaard, D.E.; Hsu, D.D. Techno-Economic 

Analysis of Biomass Fast Pyrolysis to Transportation Fuels. NREL/TP-6A2-46586. 

Golden, CO: National Renewable Energy Laboratory, 2010. 

3. Frey, H.C.; Akunuri, N. Probabilistic Modeling and Evaluation of the Performance, 

Emissions, and Cost of Texaco Gasifier-based Integrated Gasification Combined Cycle 

Systems using ASPEN [software]. Raleigh, NC: North Carolina State University, January 

2001. 

4. Bain, R.L. Material and Energy Balances for Methanol from Biomass Using Biomass 

Gasifiers. NREL/TP-510-17098. Golden, CO: National Renewable Energy Laboratory, 

January 1992. 

5. Brown, R.C. Biorenewable Resources: Engineering New Products from Agriculture. 

Ames IA: Iowa State University Press, 2003. 

6. de Swaan Arons, J.; van der Kooi, H.; Sankaranarayanan, K. “Biomass Production and 

Conversion.” Chapter 16, Efficiency and Sustainability in the Energy and Chemical 

Industries. New York, NY: Marcel Dekker, Inc., 2004. 

7. Cushman, J.H.; Easterly, J.C.; Erbach, D.C.; Foust, T.D.; Hess, J.R.; Hettenhaus, J.R.; 

Hoskinson, R.L.; Sheehan, J.J.; Sokhansanj, S.; Tagore, S.; Thompson, D.N.; Turhollow, 

A.; Wright, L.L. Roadmap for Agriculture Biomass Feedstock Supply in the United 

States. DOE/NE-ID-11129. Washington, DC: U.S. Department of Energy, 2003. 

8. Greene, N.; Celik, F.E.; Dale, B.; Jackson, M. Growing Energy: How Biofuels can Help 

End America's Oil Dependence. New York, NY: Natural Resources Defense Council, 

December 2004. 

9. Knoef, H. “Practical Aspects of Biomass Gasification.” Handbook Biomass Gasification. 

Knoef, H., ed. Enschede, Netherlands: BTG Biomass Technology Group, 2005. 

10. Knoef, H. “History of Biomass Gasification.” Handbook Biomass Gasification. Knoef, 

H., ed. Enschede, Netherlands: BTG Biomass Technology Group, 2005. 

11. Probstein, R.F.; Hicks, R.E. Synthetic Fuels. Mineola, NY: Dover Publications, 2006. 

12. Prins, M.J. Thermodynamic Analysis of Biomass Gasification and Torrefaction. 

Eindhoven, Netherlands: Technische Universiteit Eindhoven, 2005. 
51 

13. McKendry, P. “Energy Production from Biomass (Part 3): Gasification Technologies.” 

Bioresource Technology; Vol. 83(1), 2002; pp. 55-63. 

14. Warnecke, R. “Gasification of Biomass: Comparison of Fixed Bed and Fluidized Bed 

Gasifier.” Biomass and Bioenergy; Vol. 18(6), 2000; pp. 489-497. 

15. Ciferno, J.P.; Marano, J.J. Benchmarking Biomass Gasification Technologies for Fuels, 

Chemicals, and Hydrogen Production. Pittsburgh, PA: National Energy Technology 

Laboratory, June 2002. 

16. Williams, R.H.; Larson, E.D.; Katofsky, R.E.; Chen, J. “Methanol and Hydrogen from 

Biomass for Transportation.” Energy for Sustainable Development; Vol. 1(5), 1995; pp. 

18-34. 

17. van der Drift, H.; Boerrigter, H.; Coda, B.; Cieplik, M.K.; Hemmes, K. Entrained Flow 

Gasification of Biomass: Ash Behavior, Feeding Issues, and System Analyses. ECN-C--

04-039. Petten, Netherlands: Energy Research Center of the Netherlands, April 2004. 

18. Henrich, E.; Weirich, F. “Pressurised Entrained Flow Gasifiers for Biomass.” Annual 

International Conference on Thermal Treatment Technologies, New Orleans, Louisiana, 

2002. 

19. Henrich, E.; Raffelt, K.; Stahl, R.; Weirich, F. “Clean Syngas from Bio-oil/Char 

Slurries.” Bridgwater, A.V.; Boocock, D.G.B., eds. Science in Thermal and Chemical 

Biomass Conversion, Vol. 2. Newbury, UK: CPL Press, 2006; pp. 1565-1579. 

20. Naimi, L.J.; Sokhansanj, S.; Mani, S.; Hoque, M; Bi, T.; Womac, A.R.; Narajan, S. “Cost 

and Performance of Woody Biomass Size Reduction for Energy Production.” 

CSBE/SCGAB 2006 Annual Conference, American Society of Agricultural and Biological 

Engineers, Edmonton, Alberta, 2006. 

21. Amos, W.A. Report on Biomass Drying Technology. NREL/TP-570-25885. Golden, CO: 

National Renewable Energy Laboratory, November 1998. 

22. Zwart, R.W.R.; Boerrigter, H.; van der Drift, A. “The Impact of Biomass Pretreatment on 

the Feasibility of Overseas Biomass Conversion to Fischer-Tropsch Products.” Energy & 

Fuels; Vol. 20(5), 2006; pp. 2192-2197. 

23. Craig, K.R.; Mann, M.K. Cost and Performance Analysis of Biomass-Based Integrated 

Gasification Combined-Cycle (BIGCC) Power Systems. NREL/TP--430-21657. Golden, 

CO: National Renewable Energy Laboratory, 1996. 

24. Stevens, D.J. Hot Gas Conditioning: Recent Progress with Larger-scale Biomass 

Gasification Systems. NREL/SR-510-29952. Work performed by Pacific Northwest 

National Laboratory, Richland, WA. Golden, CO: National Renewable Energy 

Laboratory, August 2001. 
52 

25. Milne, T.A.; Evans, R.J.; Abatzoglou, N. Biomass Gasifier ‘Tars’: Their Nature, 

Formation, and Conversion. NREL/TP-570-25357. Golden, CO: National Renewable 

Energy Laboratory, November 1998. 

26. Stevens, D.J. Review and Analysis of the 1980-1989 Biomass Thermochemical 

Conversion Program. NREL/TP-421-7501. Golden, CO: National Renewable Energy 

Laboratory, September 1994. 

27. Nexant, Inc. Equipment Design and Cost Estimation for Small Modular Biomass Systems, 

Synthesis Gas Cleanup, and Oxygen Separation Equipment. NREL/SR-510-39945. Work 

performed by Nexant, Inc., San Francisco, CA. Golden, CO: National Renewable Energy 

Laboratory, May 2006. 

28. Wender, I. “Reactions of Synthesis Gas.” Fuel Processing Technology; Vol. 48(3), 1996; 

pp. 189-297. 

29. Consonni, S.; Larson, E.D. “Biomass-Gasifier/Aeroderivative Gas Turbine Combined 

Cycles: Part A - Technologies and Performance Modeling.” Journal of Engineering of 

Gas Turbines and Power; Vol. 118(3), 1996; pp. 507-515. 

30. Bridgwater, A.V.; Toft, A.J.; Brammer, J.G. “A Techno-Economic Comparison of Power 

Production by Biomass Fast Pyrolysis with Gasification and Combustion.” Renewable & 

Sustainable Energy Reviews; Vol. 6(3), 2002; pp. 181-246. 

31. Spath, P.L.; Dayton, D.C. Preliminary Screening -- Technical and Economic Assessment 

of Synthesis Gas to Fuels and Chemicals with Emphasis on the Potential for BiomassDerived Syngas. NREL/TP-510-34929. Golden, CO: National Renewable Energy 

Laboratory, December 2003. 

32. Huber, G.W.; Iborra, S.; Corma, A. “Synthesis of Transportation Fuels from Biomass: 

Chemistry, Catalysts, and Engineering.” Chemical Reviews; Vol. 106(9), 2006; pp. 4044-

4098. 

33. Cocco, D.; Pettinau, A.; Cau, G. “Energy and Economic Assessment of IGCC Power 

Plants Integrated with DME Synthesis Process.” Proceedings of the Institution of 

Mechanical Engineers - Part A - Power & Energy; Vol. 220(2), 2006; pp. 95-102. 

34. Dry, M. “The Fischer-Tropsch Process: 1950-2000.” Catalysis Today; Vol. 71(3-4), 

2002. 

35. Espinoza, R.; Steynberg, A.; Jager, B.; Vosloo, A. “Low Temperature Fischer-Tropsch 

Synthesis from a Sasol Perspective.” Applied Catalysis A: General; Vol. 186(1), 1999; 

pp. 13-26. 

36. Dry, M.E. “Practical and Theoretical Aspects of the Catalytic Fischer-Tropsch Process.” 

Applied Catalysis A: General; Vol. 138(2), 1996; pp. 319-344. 
53 

37. Hamelinck, C.N.; Faaij, A.P.C.; den Uil, H.; Boerrigter, H. “Production of FT 

Transportation Fuels from Biomass: Technical Options, Process Analysis and 

Optimisation, and Development Potential.” Energy; Vol. 29(11), 2004; pp. 1743-1771. 

38. Larson, E.D.; Jin, H.; Celik, F.E. “Large-Scale Gasification-Based Coproduction of Fuels 

and Electricity from Switchgrass.” Biofuels, Bioproducts, and Biorefining; Vol. 3(2), 

2009; pp. 174-194. 

39. Phillips, S.; Aden, A.; Jechura, J.; Dayton, D.C.; Eggeman, T. Thermochemical Ethanol 

via Indirect Gasification and Mixed Alcohol Synthesis of Lignocellulosic Biomass. 

NREL/TP-510-41168. Golden, CO: National Renewable Energy Laboratory, April 2007. 

40. Tijmensen, M.J.A.; Faaij, A.P.C.; Hamelinck, C.N., van Hardeveld, M.R.M. “Exploration 

of the Possibilities for Production of Fischer Tropsch Liquids and Power via Biomass 

Gasification.” Biomass and Bioenergy; Vol. 23(2), 2002; pp. 129-152. 

41. Lau, F.S.; Bowen, D.A.; Dihu, R.; Doong, S.; Hughes, E.E.; Remick, R.; Slimane, R.; 

Turn, S.Q.; Zabransky, R. Techno-Economic Analysis of Hydrogen Production by 

Gasification of Biomass. DE-FC36-01GO11089. Des Plaines, IL: Gas Technology 

Institute, December 2002. 

42. Mani, S.; Tabil, L.G.; Sokhansanj, S. “Grinding Performance and Physical Properties of 

Wheat and Barley Straws, Corn Stover and Switchgrass.” Biomass and Bioenergy; Vol. 

27(4), 2004; pp. 339-352. 

43. Higman, C.; van der Burgt, M.J. Gasification. Boston, MA: Elsevier/Gulf Professional 

Publishing, 2003. 

44. Gas Processors Suppliers Association. Engineering Data Book. 12th ed., FPS; Tulsa, OK: 

Gas Processors Suppliers Association, 2004. 

45. Merichem Company. “LO-CAT H2S Removal.” http://www.gtpmerichem.com/products/lo-cat/process.php (Accessed June 20, 2008). 

46. Kohl, A.L.; Nelson, R.B. Gas Purification. 5th ed.; Houston, TX: Gulf Professional 

Publishing, 1997. 

47. Imperial Chemical Industries AD. Catalyst Handbook: With Special Reference to Unit 

Processes in Ammonia and Hydrogen Manufacture. London, UK: Wolfe, 1970. 

48. Song, H.; Ramkrishna, D.; Trinh, S.; Wright, H. “Operating Strategies for FischerTropsch Reactors: A Model-Directed Study.” Korean Journal of Chemical Engineering; 

Vol. 21(2), 2004; pp. 308-317. 

49. Shah, P.P.; Sturtevent, G.C.; Gregor, J.H.; Humbach, M.J.; Padrta, F.G.; Steigleder, K.Z. 

Fischer-Tropsch Wax Characterization and Upgrading. DOE/PC/80017-T1. Des Plaines, 

IL: UOP, Inc., 1998. 
54 

50. Peters, M.S.; Timmerhaus, K.D.; West, R.E. Plant Design and Economics for Chemical 

Engineers. New York, NY: McGraw-Hill, 2003. 

51. Reed, M.E.; van Bibber, L.; Shuster, E.; Haslbeck, J.; Rutkowski, M.; Olsen, S.; Kramer, 

S. Baseline Technical and Economic Assessment of a Commercial Scale Fischer-Tropsch 

Liquids Facility. DOE/NETL-2007/1260. Pittsburgh, PA: National Energy Technology 

Laboratory, 2007. 

52. “Chemical Engineering Plant Cost Index.” Chemical Engineering; 2008. 

53. Robinson, P.R.; Dolbear, G.E. “Commercial Hydrotreating and Hydrocracking.” 

Ancheyta, J.; Speight, J.G., eds. Hydroprocessing of Heavy Oils and Residua. Boca 

Raton, FL: CRC Press, 2007. 

54. Energy Information Administration. “U.S. Natural Gas Wellhead Price.” 

http://tonto.eia.doe.gov/dnav/ng/hist/n9190us3m.htm (Accessed March 1 2009). 

55. United States Bureau of Labor Statistics. “Labor Index,” 2008. 

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

Lignocellulosic Biomass to Ethanol Process Design and Economics Utilizing Co-current 

Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. Golden, CO: 

National Renewable Energy Laboratory, June 2002. 

57. Couper, J.R. Process Engineering Economics. New York, NY: Marcel Dekker, 2003. 

58. CE IGCC Repowering Project Bins and Lockhoppers: Topical Report. Work performed 

under contract DE-FC21-90MC26308. Combustion Engineering, Inc., 1993. 

59. Merrow, E.W.; Phillips, K.E.; Myers, C.W. Understanding Cost Growth and 

Performance Shortfalls in Pioneer Process Plants. Work performed under contract no. 

DE-AC01-79PE70078. Santa Monica, CA: Rand Corporation, 1981. 
55 

Appendix A. Techno-Economic Model Assumptions 

Financial Assumptions 

• Capital investment 

o Equity: 100% 

o Working capital (% of FCI): 15% 

• Depreciation model 

o Zero Salvage Value for both general plant and steam/ power plant 

o Type of depreciation: Double-Declining-Balance Depreciation Method (DDB) as 

per IRS Modified Accelerated Cost Recovery System (MACRS) guidelines 

 Depreciation period (years): 

• General plant: 7 

• Steam/power system: 20 

• Construction & startup: 

o Construction period (years): 2.5 

 % spent in year “-3”: 8% 

 % spent in year “-2”: 60% 

 % spent in year “-1”: 32% 

o Startup time (years): 0.5 

 Revenues (% of normal): 50% 

 Variable costs (% of normal): 75% 

 Fixed cost (% of normal): 100% 

• Other 

o Internal rate of return: 10% 

o Income tax rate: 39% 

o Operating hours per year: 8,406 

Capital Costs 

• Cost year for analysis: 2007; cost escalation is applied using the Chemical Engineering 

Plant Cost Index 

• The plant is designed based on the state of the technology, at the nth plant level of 

experience 

• Most equipment installation factors are applied using Peters et al. for solid-fluid plants 

(i.e., 3.02 installation factor) 
56 

• Materials of construction are carbon steel, stainless steel, alloys, and refractory where 

necessary 

• Sensitivity parameters involving changes in equipment size or capacity use scaling 

exponents available in literature 

Operating Costs 

• Working capital is assumed to be 15% of the total capital investment 

• Annual maintenance materials are 2% of the total installed equipment cost 

• Boiler feedwater and wastewater treatment costs are derived from prior NREL work 

• Fresh cooling water and steam costs are calculated at 10% of the required circulation rate, 

meaning a 9:1 ratio of water recycling 

• Employee salary estimation is same as that chosen by Phillips et al. 

• Employee salaries are indexed to the year of 2007 following the data of the Bureau of 

Labor Statistics 

Feedstock, Products, and By-Products 

• Feedstock is corn stover (comprising stalks, leaves, cobs, and husks) 

o Moisture content in the feedstock is 25% 

• Feed rate is 2,000 dry metric ton per day 

o The feedstock delivery logistics are not considered 

o The feedstock is delivered to the feed handling area of the plant 

• Feed cost is assumed to be $75/dry short ton at the gate 

• Gasoline and diesel products are sold for over-the-fence prices 

• Gasoline energy content is 115,000 BTU/gallon 

• Fly ash and slag incur a solids waste disposal cost 

• Solid sulfur and electricity are sold as by-products 

Process Assumptions 

For both scenarios, most of the process was modeled with the aid of Aspen Plus software. The 

process was divided by logical process areas, which are named below. 

Area 100 - Preprocessing 

• Biomass is dried down to 10% 

o Steam raised from hot flue gas is used to dry the feedstock 

o Steam to moisture removal ratio is set at 9:1 in accordance with Amos 

o Heat is provided by combusting char and unreacted syngas 

• Grinder reduces biomass to 6-mm or less
57 

o The energy required for grinding is calculated separately using literature 

correlations by Mani et al. 

Area 200 - Gasification 

• Scenario 1: Entrained-flow gasifier is modeled using thermodynamic equilibrium 

• Scenario 2: Fluidized-bed gasifier is modeled using a mass balance calculation 

• 95% purity oxygen produced from air separation unit provides oxidizer 

• Carbon dioxide is used as solids pressurization gas 

• All char produced in LT scenario is combusted for process heat 

Area 300 - Syngas Cleaning 

• Particulates, tar, and partial ammonia removal via wet scrubbing 

o Scrubbing water is recycled at 90% rate 

o Particulate handling (not modeled) 

 High-temperature gasifier: particulate decant slurry is sent back into 

slagging gasifier 

 Low-temperature gasifier: particulate decant slurry is piled and landfilled; 

excess water is sent to aerobic water treatment (not modeled) 

o Makeup water compensates for water lost via particulate slurry 

 Process water condensate is used as makeup water 

• Sour water-gas-shift occurs at equilibrium and is modeled as such 

• Carbon dioxide, hydrogen sulfide, and excess ammonia removal via amine scrubbing 

acid gas removal (AGR) at pressure 

o 99% of sulfur is removed and 90% of carbon dioxide 

o Monoethanolamine (MEA) is the scrubbing solvent 

o Carbon dioxide is vented following LO-CAT removal of H2S 

• Hydrogen sulfide is converted to solid sulfur via LO-CAT oxidation (99% conversion) 

• Ammonia can be disposed of by decomposition (not modeled) in 

o Gasifier burner (slagging gasifier) 

o Char and syngas combustor (fluidized-bed gasifier) 

• Zinc oxide and activated carbon guard bed polishing assumed (not modeled in detail) 

Area 400 - Fuel Synthesis 
58 

• Water-gas-shift occurs at equilibrium and is modeled as such 

• Pressure swing adsorption (PSA) is employed to remove excess H2 at an efficiency of 

85% and 99% purity 

o The PSA system employs two trains with six reactors each to account for all 

stages of pressurization, depressurization, purging, etc. 

o PSA adsorbers are filled 2/3 with activated carbon and 1/3 with molecular sieve 

• Syngas is catalytically converted to fuels by one-step Fischer-Tropsch synthesis followed 

by wax hydrocracking and fuel separation 

o FT synthesis employs cobalt catalyst 

o 40% syngas conversion to fuels 

o Part of the unconverted syngas is recycled 

 A fraction of the recycle is sent to the AGR to prevent CO2 buildup 

 The overall recycle ratio is about 1.9 

• A syngas purge is used as fuel in the combustor side of the biomass dryer (only in HT 

scenario) 

• Excess syngas is sent to a gas turbine for power production 

Areas 500, 600, 700 

• Hydroprocessing and product distillation costs are estimated as a “black box” based on 

literature capital cost and operating cost information from Robinson et al. 

o Literature yield data is used for estimating the relative yields of gasoline and 

diesel 

Miscellaneous 

• Combustion occurs with 20% excess oxygen 
 

59 

Appendix B. Detailed Costs 

Cost Summaries 

Figure B-1. Economic analysis summary for HT scenario 

 

Product Value ($/gal) $4.26

Total Production at Operating Capacity (MM gal / year) 41.7

Product Yield (gal / Dry US Ton Feedstock) 61.0

Delivered Feedstock Cost $/Dry US Ton $75

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal product)

 Area 100: Pretreatment $22,700,000 7% Feedstock 123.0 28.9%

 Area 200: Gasification $67,800,000 22% Steam 6.4 1.5%

 Area 300: Syngas Cleaning $33,500,000 11% Cooling Water 5.5 1.3%

 Area 400: Fuel Synthesis $49,400,000 16% Other Raw Materials 3.4 0.8%

 Area 500: Hydrocracking/Hydrotreating $33,000,000 11% Waste Disposal 1.3 0.3%

 Area 600: Power Generation $45,600,000 15% Hydroprocessing 10.6 2.5%

 Area 700: Air Separation $24,300,000 8% Fixed Costs 34.4 8.1%

 Balance of Plant $33,100,000 11% Co-product credits -13.3 -3.1%

Capital Depreciation 63.0 14.8%

Total Installed Equipment Cost $309,400,000 Average Income Tax 52.4 12.3%

Average Return on Investment 139.5 32.7%

Indirect Costs 129,700,000

 (% of TPI) 21.4% Operating Costs ($/yr)

 Project Contingency 79,000,000 Feedstock $51,300,000

Steam $2,700,000

Total Project Investment (TPI) $605,900,000 Cooling Water $2,300,000

Other Raw Matl. Costs $1,400,000

Installed Equipment Cost per Annual Gallon $7.42 Waste Disposal $1,500,000

Total Project Investment per Annual Gallon $14.52 Hydroprocessing $4,400,000

Fixed Costs $14,300,000

Loan Rate N/A Co-product credits -$5,600,000

Term (years) N/A Capital Depreciation $26,300,000

Capital Charge Factor 0.176 Average Income Tax $21,900,000

Average Return on Investment $58,200,000

Gasifier Efficiency - HHV % 82.1

Gasifier Efficiency - LHV % 87.9 Total Plant Electricity Usage (KW) 22,065

Overall Plant Efficiency (incl. electricity) - HHV % 52.7 Electricity Produced Onsite (KW) 35,880

Overall Plant Efficiency - LHV % 53.0 Electricity Purchased from Grid (KW) 0

 Electricity Sold to Grid (KW) 13,815

Availability (%) 85.0%

Plant Hours per year 7446 Plant Electricity Use (KWh/gal product) 6.1

HT Biomass-to-Liquids Scenario Summary

2,000 Dry Metric Tonnes Biomass per Day

All Currency in 2007$ and Volume in Gallons Gasoline Equivalent (GGE)

High Temperature Entrained Flow Gasifier, Sulfur Removal, Fischer-Tropsch Synthesis, Hydroprocessing, Combined Cycle Power
 

60 

Figure B-2. Economic analysis summary for LT scenario 

Product Value ($/gal) $4.83

Total Production at Operating Capacity (MM gal / year) 32.3

Product Yield (gal / Dry US Ton Feedstock) 47.2

Delivered Feedstock Cost $/Dry US Ton $75

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/gal product)

 Area 100: Pretreatment $22,700,000 9% Feedstock 158.9 32.9%

 Area 200: Gasification $28,200,000 11% Steam 10.9 2.2%

 Area 300: Syngas Cleaning $29,300,000 12% Cooling Water 7.8 1.6%

 Area 400: Fuel Synthesis $58,700,000 23% Other Raw Materials 4.1 0.8%

 Area 500: Hydrocracking/Hydrotreating $29,500,000 12% Waste Disposal 1.5 0.3%

 Area 600: Power Generation $38,900,000 15% Hydroprocessing 9.4 2.0%

 Area 700: Air Separation $19,500,000 8% Fixed Costs 38.4 8.0%

 Balance of Plant $27,200,000 11% Co-product credits -20.4 -4.2%

Capital Depreciation 67.2 13.9%

Total Installed Equipment Cost $253,900,000 Average Income Tax 55.9 11.6%

Average Return on Investment 149.5 31.0%

Indirect Costs 107,200,000

 (% of TPI) 21.5% Operating Costs ($/yr)

 Project Contingency 65,000,000 Feedstock $51,300,000

Steam $3,500,000

Total Project Investment (TPI) $498,300,000 Cooling Water $3,500,000

Other Raw Matl. Costs $1,300,000

Installed Equipment Cost per Annual Gallon $7.86 Waste Disposal $1,500,000

Total Project Investment per Annual Gallon $15.43 Hydroprocessing $3,000,000

Fixed Costs $12,400,000

Loan Rate N/A Co-product credits -$6,600,000

Term (years) N/A Capital Depreciation $21,700,000

Capital Charge Factor 0.177 Average Income Tax $18,000,000

Average Return on Investment $48,300,000

Gasifier Efficiency - HHV % 64.3

Gasifier Efficiency - LHV % 68.8 Total Plant Electricity Usage (KW) 15,044

Overall Plant Efficiency - HHV % 43.0 Electricity Produced Onsite (KW) 31,420

Overall Plant Efficiency - LHV % 43.3 Electricity Purchased from Grid (KW) 0

 Electricity Sold to Grid (KW) 16,376

Availability (%) 85.0%

Plant Hours per year 7446 Plant Electricity Use (KWh/gal product) 5.4

LT Biomass-to-Liquids Process Engineering Analysis

2,000 Dry Metric Tonnes Biomass per Day

All Currency in 2007$ and Volume in Gallons Gasoline Equivalent (GGE)

Low Temperature Fluidized Gasifier, Sulfur Removal, Fischer-Tropsch Synthesis, Hydroprocessing, Combined Cycle Power
61 

Detailed Equipment Lists High-Temperature Scenario Equipment Lists 

Table B-1. Detailed Equipment List for Areas 100 and 200 of HT Scenario 

Equipment Number

Number 

Required

Number 

Spares Equipment Name

Original Equip Cost 

(per unit) Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaled Uninstalled 

Cost in 2007$ Installed Cost Base Year Installed Cost in 

A100.CONV1 2 Bale Transport Conveyor $400,000 2000 $800,000 $1,066,531 $1,296,000 $

A100.CONV2 2 Bale Unwrapping Conveyor $150,000 2000 $300,000 $399,949 $357,000 A100.CONV3 1 Belt Press Discharge Conveyor $50,000 2000 $50,000 $66,658 $94,500 A100.SCALE 2 Truck Scales $34,000 2000 $68,000 $90,655 $167,960 A100.FORK1 4 1 Truck Unloading Forklift $18,000 2000 $90,000 $119,985 $90,000 A100.FORK2 4 Bale Moving Forklift $18,000 2000 $72,000 $95,988 $72,000 

A100.SLAB 1 Concrete Feedstock-Storage Slab $450,655 2000 $450,655 $600,797 $991,441 $

A100.MAGSEP 1 Magnetic Separator $13,863 1998 $13,863 $18,700 $18,022 

A100.A100CHOP.CHGRIN01 4 Chopper $105,100 2007 $420,400 $420,400 $1,105,258 $

A100.A100CHOP.CHMIX01 1 Chopper Conveyor $61,400 2007 $61,400 $61,400 $185,428 A100.A100CHOP.CHSEP01 1 Chopper Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816 A100.A100DRY.DRDRY01 10 Dryer $633,700 2007 $6,337,000 $6,337,000 $15,201,647 $1A100.A100GRIN.GRGRIN01 4 Grinder $167,100 2007 $668,400 $668,400 $1,757,266 $

A100.A100GRIN.GRMIX01 1 Grinder Conveyor $61,400 2007 $61,400 $61,400 $185,428 A100.A100GRIN.GRSEP01 1 Grinder Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816 A100 Subtotal $9,434,718 $10,049,464 $21,647,582 $2A200.A200COMB.CBREAC01 1 Combustor - Steam Boiler $1,450,500 2007 $1,450,500 $1,450,500 $4,380,510 $

A200.A200SLAG.SLREAC01 1 Entrained Flow, Slagging Gasifier $23,234,043 2006 $23,234,043 $24,433,879 $54,600,000 $5

A200.A200SLAG.SLSEP01 1 Slag collector/separator $35,100 2007 $35,100 $35,100 $106,002 A200.A200SLAG.SLSEP03 3 Direct Quench Syngas Cooler $396,200 2007 $1,188,600 $1,188,600 $3,589,572 $

A200.GSHOP01 1 Biomass Receiving Hopper $151,400 2007 $297,900 $297,900 $899,658 A200.GSTANK01 1 Lockhopper $229,100 2007 $229,100 $229,100 $691,882 A200.GSTANK02 1 Biomass Feeding Bin $228,900 2007 $228,900 $228,900 $691,278 A200 Subtotal $26,664,143 $27,863,979 $64,958,902 $6
62 

Table B-2. Detailed Equipment List for Areas 300, 400, and 500 of HT Scenario 

Equipment Number

Number 

Required

Number 

Spares Equipment Name

Original Equip Cost 

(per unit) Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaled Uninstalled 

Cost in 2007$ Installed Cost Base Year Installed Cost in 

A300.A300AGR.AGRarea 1 High Pressure Amine System $6,949,800 2005 $6,949,800 $7,798,857 $20,988,396 $

A300.A300SGS.SGCOMP01 2 Sour Water Gas Shift Steam Compressor $1,381,900 2007 $2,763,800 $2,763,800 $3,316,560 

A300.A300SGS.SGREAC01 1 Sour Water Gas Shift Reactor $66,600 2007 $66,600 $66,600 $201,132 

A300.A300SUL.SUCOL01 1 LO-CAT Absorber $23,800 2007 $23,800 $23,800 $71,876 A300.A300SUL.SUREAC01 1 LO-CAT Oxidizer Vessel $1,000,000 2007 $1,000,000 $1,000,000 $3,020,000 A300.A300SUL.SUSEP01 1 Sulfur Separator $15,900 2007 $15,900 $15,900 $48,018 A300.CLCOMP01 2 Carbon Dioxide Compressor $1,181,200 2007 $2,362,400 $2,362,400 $2,834,880 

A300.CLDRUM01 1 Liquid Collection Tank $29,600 2007 $29,600 $29,600 $89,392 A300.CLHEAT03 1 Direct Quench Syngas Cooler $91,500 2007 $91,500 $91,500 $276,330 

A300.CLMIX01 1 Venturi Scrubber $27,100 2007 $27,100 $27,100 $81,842 A300 Subtotal $13,330,500 $14,179,557 $30,928,426 $3

A400.FSCOMP01 2 Booster Syngas Compressor $1,007,100 2007 $2,014,200 $2,014,200 $2,417,040 A400.FSCOMP02 1 Recycle Syngas Booster Compressor $748,400 2007 $748,400 $748,400 $898,080 

A400.FSCOMP03 1 PSA Booster Compressor $1,461,700 2007 $1,461,700 $1,461,700 $1,754,040 A400.FSHEAT01 1 Syngas Heater $73,400 2007 $73,400 $73,400 $221,668 

A400.FSHEAT03 1 Syngas Cooler $137,400 2007 $137,400 $137,400 $414,948 

A400.FSHEAT04 1 Recycle Syngas Pre-heater $21,500 2007 $21,500 $21,500 $64,930 

A400.FSREAC01 1 Fischer-Tropsch Reactor $8,888,889 2003 $8,888,889 $11,617,468 $32,000,000 $

A400.FSSEP01 2 ZnO Sulfur Removal Beds $61,000 2007 $122,000 $122,000 $368,440 

A400.FSSEP02 12 Pressure Swing Absorption Unit $33,300 2007 $399,600 $399,600 $1,206,792 A400.FSSEP03 1 FT knock-out Column $39,600 2007 $39,600 $39,600 $119,592 

A400.FSSEP04 1 Water Separator $47,900 2007 $47,900 $47,900 $144,658 

A400 Subtotal $13,954,589 $16,683,168 $39,610,188 $4

A500.HYREAC01 1 Hydroprocessing Unit $9,377,483 2007 $9,377,483 $9,377,483 $28,320,000 $

A500.HYTANK01 1 Diesel 30-day Storage Tank $1,167,600 2007 $1,167,600 $1,167,600 $3,526,152 

A500.HYTANK02 1 Gasoline 30-day Storage Tank $371,900 2007 $371,900 $371,900 $1,123,138 

A500 Subtotal $10,916,983 $10,916,983 $32,969,290 $3
63 

Table B-3. Detailed Equipment List for Areas 600 and 700 of HT Scenario 

Equipment Number

Number 

Required

Number 

Spares Equipment Name

Original Equip Cost 

(per unit) Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaled Uninstalled 

Cost in 2007$ Installed Cost Base Year Installed Cost in 

A600.COMBB 1 Combustion Turbine - Electric Generator $22,404,000 2007 $22,404,000 $22,404,000 $26,884,800 $

A600.CWPUMP 1 1 Cooling Water Pump $5,900 2007 $11,800 $11,800 $35,636 A600.ECON1_HRSG 1 Heat Recovery Steam Generator $202,200 2007 $202,200 $202,200 $610,644 

A600.HPPUMP 1 1 High Pressure Steam Pump $266,700 2007 $533,400 $533,400 $1,610,868 

A600.HPSEP 1 High Pressure Steam/Water Separation $107,400 2007 $107,400 $107,400 $324,348 

A600.LPEXP_ELECGEN 1 Combined Steam Turbine - Electric Gen. $4,709,600 2007 $4,709,600 $4,709,600 $5,651,520 

A600.LPSEP 1 Low Pressure Water/Steam Separation $108,800 2007 $108,800 $108,800 $328,576 

A600.O2COMP 1 Air Compressor $8,431,900 2007 $8,431,900 $8,431,900 $10,118,280 $

A600 Subtotal $36,509,100 $36,509,100 $45,564,672 $4

A700.COMP1 2 Air Compressor $3,346,500 2007 $6,693,000 $6,693,000 $8,031,600 

A700.COOLER 1 Air Cooler $27,200 2007 $27,200 $27,200 $82,144 A700.GOXCLR-1 1 Oxygen Compressor Cooler $23,300 2007 $23,300 $23,300 $70,366 A700.GOXCLR-2 1 Oxygen Compressor Cooler $23,000 2007 $23,000 $23,000 $69,460 A700.GOXCMP-1 2 Oxygen Compressor $1,489,600 2007 $2,979,200 $2,979,200 $3,575,040 

A700.HIGH-P.cond 1 High Pressure Column Condenser $20,300 2007 $20,300 $20,300 $61,306 A700.HIGH-P.cond acc 1 High Pressure Column Condenser Accumulator $40,500 2007 $40,500 $40,500 $122,310 

A700.HIGH-P.reflux pump 1 1 High Pressure Column Reflux Pump $14,300 2007 $28,600 $28,600 $86,372 A700.HIGH-P.tower 1 High Pressure Column Tower $314,300 2007 $314,300 $314,300 $949,186 

A700.INTRC1 1 Air Compressor Intercooler $338,300 2007 $338,300 $338,300 $1,021,666 

A700.INTRC2 1 Air Compressor Intercooler $304,500 2007 $304,500 $304,500 $919,590 

A700.INTRC3 1 Air Compressor Intercooler $222,500 2007 $222,500 $222,500 $671,950 

A700.LOW-P.reb 1 Low Pressure Column Reboiler $19,600 2007 $19,600 $19,600 $59,192 A700.LOW-P.tower 1 Low Pressure Column Tower $2,581,600 2007 $2,581,600 $2,581,600 $7,796,432 

A700.TSA 1 Water Knock-out Drum $35,900 2007 $35,900 $35,900 $108,418 

A700.TURB-1 2 Gas Expander $86,100 2007 $172,200 $172,200 $520,044 

A700.WK01 1 Water Knock-out Drum $57,700 2007 $57,700 $57,700 $174,254 

A700 Subtotal $13,881,700 $13,881,700 $24,319,330 $2

$124,691,733 $130,083,951 $259,998,390 $27$139,654,741 $145,694,026 $291,198,196 $30

TotalTotal (with BOP)
64 

Low-Temperature Scenario Equipment Lists 

Table B-4. Detailed Equipment List for Areas 100 and 200 of LT Scenario 

Equipment Number

Number 

Required

Number 

Spares Equipment Name

Original Equip Cost 

(per unit) in Base Year Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaled Uninstalled 

Cost in 2007$ Installed Cost Base Year Installed Cost in 

A100.CONV1 2 Bale Transport Conveyor $400,000 2000 $800,000 $1,066,531 $1,296,000 $A100.CONV2 2 Bale Unwrapping Conveyor $150,000 2000 $300,000 $399,949 $357,000 

A100.CONV3 1 Belt Press Discharge Conveyor $50,000 2000 $50,000 $66,658 $94,500 

A100.SCALE 2 Truck Scales $34,000 2000 $68,000 $90,655 $167,960 

A100.FORK1 4 1 Truck Unloading Forklift $18,000 2000 $90,000 $119,985 $90,000 

A100.FORK2 4 Bale Moving Forklift $18,000 2000 $72,000 $95,988 $72,000 

A100.SLAB 1 Concrete Feedstock-Storage Slab $450,655 2000 $450,655 $600,797 $991,441 $A100.MAGSEP 1 Magnetic Separator $13,863 1998 $13,863 $18,700 $18,022 

A100.A100CHOP.CHGRIN01 4 Chopper $105,100 2007 $420,400 $420,400 $1,105,258 $

A100.A100CHOP.CHMIX01 1 Chopper Conveyor $61,400 2007 $61,400 $61,400 $185,428 A100.A100CHOP.CHSEP01 1 Chopper Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816 

A100.A100DRY.DRDRY01 10 Dryer $633,700 2007 $6,337,000 $6,337,000 $15,201,647 $A100.A100GRIN.GRGRIN01 4 Grinder $167,100 2007 $668,400 $668,400 $1,757,266 $

A100.A100GRIN.GRMIX01 1 Grinder Conveyor $61,400 2007 $61,400 $61,400 $185,428 A100.A100GRIN.GRSEP01 1 Grinder Screen with Recycle Conveyor $20,800 2007 $20,800 $20,800 $62,816 

A100 Subtotal $9,434,718 $10,049,464 $21,647,582 $2

A200.A200COMB.CBCYC01 3 Combustor Cyclone (medium efficiency) $35,400 2007 $106,200 $106,200 $320,724 A200.A200COMB.CBCYC02 3 Combustor Cyclone (high efficiency) $6,700 2007 $20,100 $20,100 $60,702 

A200.A200COMB.CBMIX01 1 Ash Storage Vessel $142,800 2007 $142,800 $142,800 $431,256 A200.A200COMB.CBREAC01 1 Combustor - Steam Boiler $1,450,500 2007 $1,450,500 $1,450,500 $4,380,510 $

A200.A200CYC.CYCYC01 2 1st train, medium efficiency cyclone $20,300 2007 $40,600 $40,600 $122,612 A200.A200CYC.CYCYC02 4 1st train, high efficiency cyclone $24,900 2007 $99,600 $99,600 $300,792 A200.A200CYC.CYMIX02 1 Char Collector and conveyor $84,400 2007 $84,400 $84,400 $254,888 A200.GSREAC01 7 Fluidized Bed Gasifier (Pressurized) $1,096,170 2003 $7,673,191 $10,028,594 $14,843,424 $A200.GSTANK01 7 Biomass Receiving Hopper $71,700 2007 $501,900 $501,900 $1,247,712 $

A200.GSTANK02 7 Lockhopper $47,700 2007 $333,900 $333,900 $830,068 A200.GSTANK03 7 Biomass Feeding Bin $47,700 2007 $333,900 $333,900 $830,068 

A200 Subtotal $10,787,091 $13,142,494 $23,622,756 $2
65 

Table B-5. Detailed Equipment List for Areas 300, 400, and 500 of LT Scenario 

Equipment Number

Number 

Required

Number 

Spares Equipment Name

Original Equip Cost 

(per unit) in Base Year Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaled Uninstalled 

Cost in 2007$ Installed Cost Base Year Installed Cost in 2007

A300.A300AGR.AGRarea 1 High Pressure Amine System $6,050,000 2005 $6,050,000 $6,789,129 $18,271,000 $20,50A300.A300SUL.SUCOL01 1 LO-CAT Absorber $16,200 2007 $16,200 $16,200 $48,924 $4A300.A300SUL.SUREAC01 1 LO-CAT Oxidizer Vessel $1,000,000 2007 $1,000,000 $1,000,000 $3,020,000 $3,02

A300.A300SUL.SUSEP01 1 Sulfur Separator $16,200 2007 $16,200 $16,200 $48,924 $4

A300.CLCMP01 2 Carbon Dioxide Compressor $1,176,900 2007 $2,353,800 $2,353,800 $2,824,560 $2,82

A300.CLHEAT01 2 Direct Quench Recycle Cooling $188,800 2007 $377,600 $377,600 $1,140,352 $1,14A300.CLHEAT02 1 Venturi Recycle Cooling $91,500 2007 $91,500 $91,500 $276,330 $27

A300.CLMIX01 1 Venturi Scrubber $26,800 2007 $26,800 $26,800 $80,936 $8A300.CLSEP03 2 Direct Quench Syngas Cooler $188,800 2007 $377,600 $377,600 $1,140,352 $1,14

A300.CLSEP04 1 Venturi Liquid Collection Tank $74,500 2007 $74,500 $74,500 $224,990 $22

A300 Subtotal $10,384,200 $11,123,329 $27,076,368 $29,308A400.A400COND.CDHEAT01 1 Syngas Heater $60,500 2007 $60,500 $60,500 $182,710 $18

A400.A400COND.CDHEAT02 1 Syngas Pre-heater Furnace $1,949,500 2007 $1,949,500 $1,949,500 $5,887,490 $5,88

A400.A400COND.CDHEAT03 1 Reformed Syngas Waste Heat Boiler $396,600 2007 $396,600 $396,600 $1,197,732 $1,19

A400.A400COND.CDHEAT04 1 Syngas Cooler #2 $41,200 2007 $41,200 $41,200 $124,424 $12

A400.A400COND.CDREAC01 1 Steam Methane Reformer $1,650,800 2007 $1,650,800 $1,650,800 $4,985,416 $4,98

A400.A400COND.CDREAC02 1 Water Gas Shift Reactor $136,600 2007 $136,600 $136,600 $412,532 $41A400.A400COND.CDSEP01 2 ZnO Sulfur Removal Beds $46,400 2007 $92,800 $92,800 $280,256 $28A400.FSCOMP01 2 Booster Syngas Compressor $921,600 2007 $1,843,200 $1,843,200 $2,211,840 $2,21

A400.FSCOMP02 1 Recycle Syngas Booster Compressor $725,400 2007 $725,400 $725,400 $870,480 $87A400.FSCOMP03 1 PSA Booster Compressor $1,482,100 2007 $1,482,100 $1,482,100 $1,778,520 $1,77A400.FSDRUM01 1 PSA Knock-out $1,482,100 2007 $1,482,100 $1,482,100 $4,475,942 $4,47

A400.FSHEAT03 1 Syngas Cooler $165,200 2007 $165,200 $165,200 $498,904 $49

A400.FSHEAT04 1 Recycle Syngas Pre-heater $24,300 2007 $24,300 $24,300 $73,386 $7

A400.FSREAC01 1 Fischer-Tropsch Reactor $7,303,889 2003 $7,303,889 $9,545,928 $26,294,000 $34,36

A400.FSSEP02 12 Pressure Swing Absorption Unit $30,500 2007 $366,000 $366,000 $1,105,320 $1,10A400.FSSEP03 1 FT knock-out Column $72,100 2007 $72,100 $72,100 $217,742 $21

A400.FSSEP04 1 Water Separator $39,200 2007 $39,200 $39,200 $118,384 $11

A400 Subtotal $17,792,289 $20,034,328 $50,596,694 $58,668A500.HYREAC01 1 Hydrocracking/Hydrotreating Unit $7,927,152 2007 $7,927,152 $7,927,152 $23,940,000 $23,94

A500.HYTANK01 1 Gasoline 30-day Storage Tank $646,300 2007 $646,300 $646,300 $1,951,826 $1,95A500.HYTANK02 1 Diesel 30-day Storage Tank $1,200,700 2007 $1,200,700 $1,200,700 $3,626,114 $3,62

A500 Subtotal $9,774,152 $9,774,152 $29,517,940 $29,517
66 

Table B-6. Detailed Equipment List for Areas 600 and 700 of LT Scenario 

Equipment Number

Number 

Required

Number 

Spares Equipment Name

Original Equip Cost 

(per unit) in Base Year Base Year

Total Original Equip Cost 

(Req'd & Spare) in Base 

Year

Scaled Uninstalled 

Cost in 2007$ Installed Cost Base Year Installed Cost in 

A600.COMBB 1 Combustion Turbine - Electric Generator $18,607,700 2007 $18,607,700 $18,607,700 $22,329,240 $2

A600.CWPUMP 1 1 Cooling Water Pump $5,900 2007 $11,800 $11,800 $35,636 

A600.ECON1_HRSG 1 Heat Recovery Steam Generator $202,200 2007 $202,200 $202,200 $610,644 A600.HPPUMP 1 1 High Pressure Steam Pump $266,700 2007 $533,400 $533,400 $1,610,868 $

A600.HPSEP 1 High Pressure Steam/Water Separation $107,400 2007 $107,400 $107,400 $324,348 A600.LPEXP_ELECGEN 1 Combined Steam Turbine - Electric Gen. $5,056,300 2007 $5,056,300 $5,056,300 $6,067,560 $

A600.LPSEP 1 Low Pressure Water/Steam Separation $108,800 2007 $108,800 $108,800 $328,576 A600.O2COMP 1 Air Compressor $6,331,200 2007 $6,331,200 $6,331,200 $7,597,440 $

A600 Subtotal $30,958,800 $30,958,800 $38,904,312 $3

A700.COMP1 2 Air Compressor $3,119,600 2007 $6,239,200 $6,239,200 $7,487,040 $

A700.COOLER 1 Air Cooler $24,300 2007 $24,300 $24,300 $73,386 

A700.GOXCLR-1 1 Oxygen Compressor Cooler $23,300 2007 $23,300 $23,300 $70,366 

A700.GOXCLR-2 1 Oxygen Compressor Cooler $23,000 2007 $23,000 $23,000 $69,460 

A700.GOXCMP-1 2 Oxygen Compressor $1,514,700 2007 $3,029,400 $3,029,400 $3,635,280 $

A700.HIGH-P.cond 1 High Pressure Column Condenser $20,300 2007 $20,300 $20,300 $61,306 

A700.HIGH-P.cond acc 1 High Pressure Column Condenser Accumulator $36,300 2007 $36,300 $36,300 $109,626 A700.HIGH-P.reflux pump 1 1 High Pressure Column Reflux Pump $14,300 2007 $28,600 $28,600 $34,320 

A700.HIGH-P.tower 1 High Pressure Column Tower $279,900 2007 $279,900 $279,900 $335,880 A700.INTRC1 1 Air Compressor Intercooler $338,300 2007 $338,300 $338,300 $405,960 A700.INTRC2 1 Air Compressor Intercooler $304,500 2007 $304,500 $304,500 $919,590 A700.INTRC3 1 Air Compressor Intercooler $222,500 2007 $222,500 $222,500 $671,950 A700.LOW-P.reb 2 Low Pressure Column Reboiler $19,600 2007 $39,200 $39,200 $118,384 A700.LOW-P.tower 1 Low Pressure Column Tower $1,538,900 2007 $1,538,900 $1,538,900 $4,647,478 $

A700.TSA 1 Water Knock-out Drum $30,100 2007 $30,100 $30,100 $90,902 

A700.TURB-1 2 Gas Expander $89,200 2007 $178,400 $178,400 $538,768 A700.WK01 1 Water Knock-out Drum $64,800 2007 $64,800 $64,800 $195,696 A700 Subtotal $12,421,000 $12,421,000 $19,465,392 $1

Total $101,552,251 $107,503,567 $210,831,043 $22Total (with BOP) $113,738,521 $120,403,995 $236,130,768 $25
67 

Discounted Cash Flow High-Temperature Scenario 

Table B-7. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of HT Scenari

DCFROR Worksheet

Year -2 -10 1 2 3 4 5 6 

Fixed Capital Investment $50,890,395 $316,115,651 $168,595,014Working Capital $79,028,913

Loan Payment $0 $0 $0 $0 $0 $0 

 Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 

 Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 

 GGE (Gallon of Gasoline Equiv.) Sales $133,364,635 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 

 Diesel Sales $0 $0 $0 $0 $0 $0 

 By-Product Credit $4,173,208 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 

Plant Performance 1.00 1.00 1.00 1.00 1.00 1.00 

Total Annual Sales $137,537,843 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 Annual Manufacturing Cost Raw Materials $44,894,145 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594  SWGS catalysts $114,621 $0 $0 $114,621 $0 $0  Steam reforming catalysts $0 $0 $0 $0 $0 $0 

 ZnO $424,410 $0 $0 $424,410 $0 $0  Pressure Swing Adsorption Packing $497,135 $0 $0 $497,135 $0 $0  FT catalysts $7,686,720 $0 $0 $7,686,720 $0 $0 

 Other Variable Costs $11,727,856 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264  Fixed Operating Costs $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 Total Product Cost $79,690,672 $79,056,643 $79,056,643 $87,779,529 $79,056,643 $79,056,643 Annual Depreciation General Plant DDB $128,361,546 $91,686,818 $65,490,585 $46,778,989 $33,413,564 $23,866,831  SL $64,180,773 $53,483,977 $45,843,409 $40,931,615 $38,982,491 $38,982,491  Remaining Value $320,903,864 $229,217,046 $163,726,461 $116,947,472 $83,533,909 $59,667,078  Actual $128,361,546 $91,686,818 $65,490,585 $46,778,989 $38,982,491 $38,982,491  Steam Plant DDB $5,819,551 $5,383,084 $4,979,353 $4,605,902 $4,260,459 $3,940,925  SL $3,879,700 $3,777,603 $3,688,410 $3,612,472 $3,550,382 $3,503,044  Remaining Value $71,774,458 $66,391,374 $61,412,021 $56,806,119 $52,545,660 $48,604,736  Actual $5,819,551 $5,383,084 $4,979,353 $4,605,902 $4,260,459 $3,940,925 Net Revenue ($76,333,925) $7,257,245 $33,857,210 $44,219,371 $61,084,198 $61,403,732 Losses Forward ($76,333,925) ($69,076,681) ($35,219,471) $0 $0 

Taxable Income ($76,333,925) ($69,076,681) ($35,219,471) $8,999,900 $61,084,198 $61,403,732 Income Tax $0 $0 $0 $3,509,961 $23,822,837 $23,947,455 Annual Cash Income $57,847,171 $104,327,147 $104,327,147 $92,094,301 $80,504,310 $80,379,692 Discount Factor 1.21 1.1 1 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 0.56447393 

Annual Present Value $645,181,377 $52,588,337 $86,220,783 $78,382,530 $62,901,646 $49,986,843 $45,372,241 Total Capital Investment + Interest $61,577,378 $347,727,216 $247,623,927Net Present Worth $0
68 

Table B-8. Discounted Cash Flow Sheet for Years 9-20 of HT Scenario 

DCFROR Worksheet

Year 9 10 11 12 13 14 15 16 17 18 Fixed Capital InvestmentWorking Capital 

Loan Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0  Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

 Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

 GGE (Gallon of Gasoline Equiv.) Sales $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 $177,819,513 

 Diesel Sales $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

 By-Product Credit $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 $5,564,277 

Plant Performance 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 

Total Annual Sales $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 $183,383,791 Annual Manufacturing Cost Raw Materials $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 

 SWGS catalysts $0 $114,621 $0 $0 $114,621 $0 $0 $114,621 $0 $0 

 Steam reforming catalysts $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

 ZnO $0 $424,410 $0 $0 $424,410 $0 $0 $424,410 $0 $0 

 Pressure Swing Adsorption Packing $0 $497,135 $0 $0 $497,135 $0 $0 $497,135 $0 $0 

 FT catalysts $0 $7,686,720 $0 $0 $7,686,720 $0 $0 $7,686,720 $0 $0 

 Other Variable Costs $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 $13,403,264 

 Fixed Operating Costs $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 $14,345,785 Total Product Cost $79,056,643 $87,779,529 $79,056,643 $79,056,643 $87,779,529 $79,056,643 $79,056,643 $87,779,529 $79,056,643 $79,056,643 Annual Depreciation General Plant DDB SL Remaining Value Actual Steam Plant

 DDB $3,119,057 $2,885,128 $2,668,743 $2,468,587 $2,283,443 $2,112,185 $1,953,771 $1,807,238 $1,671,696 $1,546,318 

 SL $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 

 Remaining Value $38,468,370 $35,583,242 $32,914,499 $30,445,912 $28,162,468 $26,050,283 $24,096,512 $22,289,273 $20,617,578 $19,071,260 

 Actual $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 $3,458,414 

Net Revenue $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 

Losses Forward $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

Taxable Income $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 $92,145,848 $100,868,733 $100,868,733 Income Tax $39,338,806 $35,936,881 $39,338,806 $39,338,806 $35,936,881 $39,338,806 $39,338,806 $35,936,881 $39,338,806 $39,338,806 Annual Cash Income $64,988,341 $59,667,381 $64,988,341 $64,988,341 $59,667,381 $64,988,341 $64,988,341 $59,667,381 $64,988,341 $64,988,341 Discount Factor 0.424097618 0.385543289 0.350493899 0.318630818 0.28966438 0.263331254 0.239392049 0.217629136 0.197844669 0.17985879 

Annual Present Value $27,561,401 $23,004,358 $22,778,017 $20,707,288 $17,283,515 $17,113,461 $15,557,692 $12,985,361 $12,857,597 $11,688,724 Total Capital Investment + Interest Net Present Worth
69 

Low-Temperature Scenario 

Table B-9. Discounted Cash Flow Sheet for Construction Period and Years 1-8 of LT Scenario 

DCFROR Worksheet

Year -2 -1 0 1 2 3 4 5 6 Fixed Capital Investment $41,888,460 $259,981,649 $138,656,880Working Capital $64,995,412Loan Payment $0 $0 $0 $0 $0 $0 

 Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0  Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0  GGE (Gallon of Gasoline Equiv.) Sales $117,025,289 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 

 Diesel Sales $0 $0 $0 $0 $0 $0  By-Product Credit $4,945,498 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 Plant Performance 1.00 1.00 1.00 1.00 1.00 1.00 Total Annual Sales $121,970,788 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 Annual Manufacturing Cost Raw Materials $44,894,145 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594  WGS catalysts $104,732 $0 $0 $104,732 $0 $0  Steam reforming catalysts $103,412 $0 $0 $103,412 $0 $0  ZnO $424,410 $0 $0 $424,410 $0 $0  Pressure Swing Adsorption Packing $497,135 $0 $0 $497,135 $0 $0  FT catalysts $6,127,680 $0 $0 $6,127,680 $0 $0  Other Variable Costs $11,238,097 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539  Fixed Operating Costs $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 Total Product Cost $75,794,444 $76,555,967 $76,555,967 $83,813,336 $76,555,967 $76,555,967 Annual Depreciation General Plant DDB $104,833,121 $74,880,801 $53,486,286 $38,204,490 $27,288,922 $19,492,087  SL $52,416,561 $43,680,467 $37,440,400 $33,428,929 $31,837,075 $31,837,075  Remaining Value $262,082,803 $187,202,002 $133,715,716 $95,511,226 $68,222,304 $48,730,217  Actual $104,833,121 $74,880,801 $53,486,286 $38,204,490 $31,837,075 $31,837,075  Steam Plant DDB $4,979,012 $4,605,586 $4,260,167 $3,940,654 $3,645,105 $3,371,722  SL $3,319,341 $3,231,990 $3,155,679 $3,090,709 $3,037,588 $2,997,087  Remaining Value $61,407,813 $56,802,227 $52,542,060 $48,601,405 $44,956,300 $41,584,577  Actual $4,979,012 $4,605,586 $4,260,167 $3,940,654 $3,645,105 $3,371,722 Net Revenue ($63,635,790) $6,585,363 $28,325,297 $36,669,236 $50,589,570 $50,862,953 Losses Forward ($63,635,790) ($57,050,426) ($28,725,129) $0 $0 Taxable Income ($63,635,790) ($57,050,426) ($28,725,129) $7,944,107 $50,589,570 $50,862,953 Income Tax $0 $0 $0 $3,098,202 $19,729,932 $19,836,551 Annual Cash Income $46,176,343 $86,071,750 $86,071,750 $75,716,179 $66,341,818 $66,235,199 Discount Factor 1.21 1.1 1 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 0.56447393 Annual Present Value $530,655,988 $41,978,494 $71,133,678 $64,666,980 $51,715,169 $41,193,049 $37,388,043 Total Capital Investment + Interest $50,685,036 $285,979,814 $203,652,292Net Present Worth $0
70 

Table B-10. Discounted Cash Flow Sheet for Years 9-20 of LT Scenario 

DCFROR Worksheet

Year 9 10 11 12 13 14 15 16 17 18 Fixed Capital InvestmentWorking Capital 

Loan Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0  Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0  Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0  GGE (Gallon of Gasoline Equiv.) Sales $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719 $156,033,719  Diesel Sales $0 $0 $0 $0 $0 $0 $0 $0 $0 $0  By-Product Credit $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 $6,593,998 Plant Performance 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 Total Annual Sales $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 $162,627,717 Annual Manufacturing Cost Raw Materials $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594 $51,307,594  WGS catalysts $0 $104,732 $0 $0 $104,732 $0 $0 $104,732 $0 $0  Steam reforming catalysts $0 $103,412 $0 $0 $103,412 $0 $0 $103,412 $0 $0  ZnO $0 $424,410 $0 $0 $424,410 $0 $0 $424,410 $0 $0  Pressure Swing Adsorption Packing $0 $497,135 $0 $0 $497,135 $0 $0 $497,135 $0 $0  FT catalysts $0 $6,127,680 $0 $0 $6,127,680 $0 $0 $6,127,680 $0 $0  Other Variable Costs $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539 $12,843,539  Fixed Operating Costs $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 $12,404,834 Total Product Cost $76,555,967 $83,813,336 $76,555,967 $76,555,967 $83,813,336 $76,555,967 $76,555,967 $83,813,336 $76,555,967 $76,555,967 Annual Depreciation General Plant DDB SL Remaining Value Actual Steam Plant  DDB $2,668,560 $2,468,418 $2,283,287 $2,112,040 $1,953,637 $1,807,115 $1,671,581 $1,546,212 $1,430,246 $1,322,978  SL $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903  Remaining Value $32,912,244 $30,443,825 $28,160,538 $26,048,498 $24,094,861 $22,287,746 $20,616,165 $19,069,953 $17,639,706 $16,316,728  Actual $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 $2,958,903 Net Revenue $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 Losses Forward $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 Taxable Income $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 $75,855,478 $83,112,848 $83,112,848 Income Tax $32,414,011 $29,583,637 $32,414,011 $32,414,011 $29,583,637 $32,414,011 $32,414,011 $29,583,637 $32,414,011 $32,414,011 Annual Cash Income $53,657,740 $49,230,744 $53,657,740 $53,657,740 $49,230,744 $53,657,740 $53,657,740 $49,230,744 $53,657,740 $53,657,740 Discount Factor 0.424097618 0.385543289 0.350493899 0.318630818 0.28966438 0.263331254 0.239392049 0.217629136 0.197844669 0.17985879 Annual Present Value $22,756,120 $18,980,583 $18,806,710 $17,097,009 $14,260,393 $14,129,760 $12,845,236 $10,714,044 $10,615,898 $9,650,816 Total Capital Investment + Interest Net Present Worth
 

71 

Appendix C. Scenario Modeling Details 

Property Method 

The model setup includes a particle size distribution in order to better estimate the solids 

simulation in the grinding and cyclone operations. It operates globally with the Redlich-KwongSoave with Boston-Mathias modification (RKS-BM) property method, which is recommended 

for medium-temperature refining and gas processing operations including combustion and 

gasification. During acid-gas absorption and stripping, another property method, ELECNRTL, is 

used for more accurate simulation. For solids handling, such as in the pretreatment area and 

cyclones, the SOLIDS property method is used. 

Stream/Block Nomenclature 

All streams and blocks within the model follow a specific alphanumeric notation with the 

purpose of clarity and consistency across scenarios and across platforms. Each area within the 

model (e.g., Area 200 gasification) has a two letter abbreviation (e.g., gasification is GS). These 

abbreviations are used for naming streams as well as blocks. In addition to serving purposes 

mentioned above, the notation is descriptive (e.g., the notation REAC describes a block as a 

reactor). Another example is SGAS, which describes a stream that contains syngas. Aspen Plus 

software limits block and stream names to be eight characters. 

Figure C-1 shows the pattern of notation for a syngas stream in the gasification area. 

Area Number Description 

G S 0 1 S G A S

Figure C-1. Stream nomenclature used in model 

Similarly, the notation for the first reactor block in the gasification area is shown in Figure C-2. 

Area Description Number

G S R E A C 0 1

Figure C-2. Block nomenclature used in model 

Table C-1 contains the abbreviations for areas, unit operation block descriptions, and stream 

descriptions. 

 
 

72 

Table C-1. Detailed Description of Stream and Block Nomenclature 

Area Description Name Block Name Stream Name 

Plant All Areas PL Reactor REAC Biomass BMAS

A100 Pretreatment PR Mixer MIX Steam STM 

A100CHOP Chopping CH Heat Mixer QMX Flue Gas FLUE 

A100DRY Drying DR Work Mixer WMX Syngas SGAS 

A100GRIN Grinding GR Splitter SPL Ash ASH 

A200 Gasification GS,SL Separator SEP Carbon Dioxide CO2

A200CYC Cyclones CY Cyclones CYC Air AIR 

A200COMB Combustion CB Flash Drum DRUM Hydrogen HYD 

A300 Syngas Cleaning CL Column COL FT Products FT 

A300AGR Acid Gas Removal AG Distillation DIST Water WAT 

A300SUL Sulfur Recovery SU Grinder GRIN Oxygen OX 

A400 Fuel Synthesis FS Dryer DRY Sulfur SUL 

A400COND Syngas Conditioning CD Heater HEAT Fuel FUEL 

A400MTG Methanol to Gasoline MG Heat Exchanger HX Tar TAR 

A500 Hydrocracking HY Tank/Hopper TANK Char CHAR 

A500 Fuel Separation SE Pump PMP Acid Gas AG 

A600 Power Generation PG Compressor COMP Lean MEA Soln. MEAL 

A700 Air Separation Unit Turbine TURB Rich MEA Soln. MEAR

 Light Gases LGAS 

 Nitrogen NTGN 

A special notation is used for heat and work streams. In the case that the first reactor in the 

gasification area includes a heat stream leaving the unit, it follows the nomenclature shown in 

Figure C-3. 

Q or W Area Block Description Number 

Q - G S R E A 1 

Figure C-3. Heat and work stream nomenclature used in model 

The Q or W sets the stream apart as a heat or work stream. The block description is limited to 

three characters and the number is limited to one character. 
 

73 

Aspen Plus Calculator Block Descriptions 

High-Temperature Scenario 

AIRCOMB 

This block calculates the nitrogen that accompanies the oxygen in the air inlet for the combustion 

of unconverted syngas. Molar nitrogen flow (in kmol/h) is calculated as follows: 

ሶܯ

ேଶ ൌ ൬

0.79

0.21൰ · ܯሶ

ைଶ (eqn. 20)

where ܯሶ

ைଶ is molar flow of oxygen in kmol/h. 

AMINE

This block calculates the molar flow of monoethanolamine (MEA) needed for the required acid 

gas removal (CO2 and H2S) arriving from syngas quench and FT unconverted syngas recycle 

stream. The MEA is able to capture 0.35 moles acid gas per mole MEA. Additionally, the MEA 

is diluted as explained in DILUTH2O. 

Molar MEA flow (in kmol/h) is calculated by 

ሶܯ

ொ஺ ൌ ሺܯሶ஼ைଶ,௦௬௡ ൅ ܯሶ஼ைଶ,௥௘௖ ൅ ܯሶுଶௌ,௦௬௡ሻ/0.35 (eqn. 21)

where ܯሶ

஼ைଶ,௦௬௡ is molar flow of CO2 from the syngas after the syngas quench, ܯሶ஼ைଶ,௥௘௖ is the 

molar flow of CO2 from the unconverted syngas recycle after the FT synthesis, and ܯሶ

ுଶௌ,௦௬௡ is 

the molar flow rate of H2S from the syngas quench. 

Because the MEA solution in the amine absorption unit is to be 20 wt% concentrated with water, 

the flow of water must be calculated. Mole flow of water is calculated as 

ሶܯ

ுଶை ൌ ܯሶ

/0.20஺ொܹܯ כ ஺ொ

ுଶைܹܯ

(eqn. 22)

BIOELEM 

Because the high-temperature gasifier is modeled at equilibrium, the simulation software 

requires that all components in the input are located in the conventional stream. Therefore, this 

block splits the biomass into the following compounds based on its ultimate analysis: carbon, 

hydrogen, oxygen, sulfur, nitrogen, and ash. Water in the biomass is not affected because it is 

already a conventional component. Biomass in the exit stream is set to zero. 
 

74 

FTDISTR 

This calculator block calculates an alpha chain growth parameter using the equation by Song et 

al. (2004) for cobalt catalyst. Inlet and outlet streams are defined and calculated. FT products 

include paraffins from C1 through C20. FT waxes are paraffins at C30. 

FT reaction is as follows: 

CO ൅ 2.1 כ Hଶ െ െ൐ െሺܥHଶሻെ ൅HଶO (eqn. 23)

Section 100 sets the CO conversion. 

Section 200 calculates the reaction extent (in lbmol) based on an alpha value of 0.9. 

 -------------Section 100------------------------ 

Percent conversion of CO is calculated as follows, and then the molar amount of converted CO 

(COCONV) is calculated knowing the molar amount of CO entering (COIN). 

PERCEN = 40 

CONV=PERCEN/100.0 

COCONV=COIN*CONV 

 ------------Section 200-------------------------- 

R1, R2, R3, etc. represent the molar reaction extent (lbmol/h) that is utilized in the FT reactor for 

each reaction (i.e., CO + 3*H2  CH4 + H2O, 2*CO + 5*H2  C2H6 + 2*H2O, etc.). The 

coefficients of each reaction extent are calculated by solving a set of 21 equations shown below 

in Table C-2 and as described in the Detailed Calculations section of this Appendix. 
 

75 

Table C-2. Reaction Extent Equations for Each Alkane Hydrocarbon 

Alkane 

Component Equation 

0.01 כ ܱܸܰܥܱܥ ൌ ܴ1 1C

0.018/2 כ ܱܸܰܥܱܥ ൌ ܴ2 2C

0.0243/3 כ ܱܸܰܥܱܥ ൌ ܴ3 3C

0.02916/4 כ ܱܸܰܥܱܥ ൌ ܴ4 4C

0.0328/5 כ ܱܸܰܥܱܥ ൌ ܴ5 5C

0.03543/6 כ ܱܸܰܥܱܥ ൌ ܴ6 6C

0.0372/7 כ ܱܸܰܥܱܥ ൌ ܴ7 7C

0.03826/8 כ ܱܸܰܥܱܥ ൌ ܴ8 8C

0.03874/9 כ ܱܸܰܥܱܥ ൌ ܴ9 9C

0.03874/10 כ ܱܸܰܥܱܥ ൌ ܴ10 10C

0.03835/11 כ ܱܸܰܥܱܥ ൌ ܴ11 11C

0.03766/12 כ ܱܸܰܥܱܥ ൌ ܴ12 12C

0.03672/13 כ ܱܸܰܥܱܥ ൌ ܴ13 13C

0.03559/14 כ ܱܸܰܥܱܥ ൌ ܴ14 14C

0.03432/15 כ ܱܸܰܥܱܥ ൌ ܴ15 15C

0.03294/16 כ ܱܸܰܥܱܥ ൌ ܴ16 16C

0.0315/17 כ ܱܸܰܥܱܥ ൌ ܴ17 17C

0.03002/18 כ ܱܸܰܥܱܥ ൌ ܴ18 18C

0.02852/19 כ ܱܸܰܥܱܥ ൌ ܴ19 19C

0.02702/20 כ ܱܸܰܥܱܥ ൌ ܴ20 20C

0.36473/30 כ ܱܸܰܥܱܥ ൌ ܴ21 30C

GRIND 

This block calculates the power requirement (in kW) for grinding the biomass from the chop size 

of 15 mm to the final size of 1 mm. This power requirement data is found in Mani et al. and is 

for 12% exiting moisture. The correlation was changed from a polynomial (quadratic) regression, 

which Mani et al. used, to a power regression because the power regression more accurately 

matched the data. ܵ௖௨௧ is the final grind size in the units of millimeters. 

ܲ௚௥௜௡ௗ ൌ ൫28.76 ڄ ܵ௖௨௧

ି଴.଼ଵ൯ڄ݉ሶ ௕௜௢௠௔௦௦ (eqn. 24)

HRSG 

This calc block totals the heat transfer areas of all the heat exchangers in A600 Power Generation 

for use in Aspen Icarus software costing of a heat recovery steam generator, which is estimated 

as a waste heat boiler. 

HUMIDITY 

This block sets humidity of the air entering the air separation unit. 
 

76 

HV-101, HV-203, HV-445 

This block calculates the lower and higher heating values of the following streams: biomass, 

syngas, and FT products. 

LOCKHOP 

This block calculates the CO2 required for pressurizing the lockhopper. Higman et al. reports 

0.09 kg of pressurization gas is required per kg of biomass. 

(25 .eqn (ܵܣܯܱܫܤ כ 0.09 ൌ ܱ2ܥܺ

MEATEMP 

This block sets the temperature of the incoming monoethanolamine solution entering the 

absorber column in the AGR area. 

MOISTURE 

This block sets the moisture content of the entering biomass to the preprocessing area and sets 

the biomass moisture content exiting the biomass dryer. Also, the steam loop flow rate for drying 

the biomass is set at 9 times the amount of moisture removed during the drying process. 

Moisture content (% wet basis) of entering biomass feed, XMOIS1 ൌ 25. Inlet mass flow of 

moisture, WATERI, is computed. 

(26 .eqn (ሻܵ1/100ܫܱܯܺ െ 1ሺܵ1/100/ܫܱܯܺ כ ܦܧܧܨ ൌ ܫܴܧܶܣܹ

Moisture content (% wet basis) of biomass exiting the dryer, XMOIS2 ൌ 10. Mass flow of 

moisture, WATERO, is computed. 

WATERO ൌ FEED כ XMOIS2/100/ሺ1 െ XMOIS2/100ሻ (eqn. 27)

Specify steam required to remove moisture, STEAMI. 

STEAMI ൌ 9 כ ሺWATERI െ WATEROሻ (eqn. 28)

O2COMB 

Oxygen is required to combust the char and syngas that provides the energy necessary for drying 

the biomass. A system of stoichiometric combustion reactions is set up to sum all the oxygen 

required to fully combust the unconverted syngas purge from the FT synthesis outlet. The 

reactions are shown in Table C-3. 
 

77 

Table C-3. Combustion Reactions to Determine Required Oxygen 

Component Reaction 

ܱ2ܥ ՜ ܱ2 · 0.5 ൅ ܱܥ CO

2ܱܪ ՜ ܱ2 · 0.5 ൅ 2ܪ 2H

ܱ2ܥ ൅ 2ܱܪ2 ՜ 2ܱ2 ൅ 4ܪܥ 4CH

ܱ2ܥ2 ൅ 2ܱܪ3 ՜ ܱ2 · 3.5 ൅ 6ܪ2ܥ 6H2C

ܱ2ܥ2 ൅ 2ܱܪ2 ՜ 3ܱ2 ൅ 4ܪ2ܥ 4H2C

ܱ2ܥ ൅ 2ܱܪ ՜ 2.5ܱ2 ൅ 2ܪ2ܥ 2H2C

ܱ2ܥ3 ൅ 2ܱܪ4 ՜ 5ܱ2 ൅ 8ܪ3ܥ 8H3C

ܱ2ܥ4 ൅ 2ܱܪ5 ՜ 6.5ܱ2 ൅ 10ܪ4ܥ 10H4C

ܱ2ܥ5 ൅ 2ܱܪ6 ՜ 8ܱ2 ൅ 12ܪ5ܥ 12H5C

ܱ2ܥ6 ൅ 2ܱܪ7 ՜ 9.5ܱ2 ൅ 14ܪ6ܥ 14H6C

ܱ2ܥ7 ൅ 2ܱܪ8 ՜ 11ܱ2 ൅ 16ܪ7ܥ 16H7C

ܱ2ܥ8 ൅ 2ܱܪ9 ՜ 12.5ܱ2 ൅ 18ܪ8ܥ 18H8C

ܱ2ܥ9 ൅ 2ܱܪ10 ՜ 14ܱ2 ൅ 20ܪ9ܥ 20H9C

ܱ2ܥ14 ൅ 2ܱܪ5 ՜ 16.5ܱ2 ൅ ሻݎܽݐሺ10ܪ14ܥ Tar

H2S ܪ2ܵ ൅ 1.5ܱ2 ՜ ܪ2ܱ ൅ ܱܵ2

ܱܰ2 ൅ 2ܱܪ1.5 ՜ 1.75ܱ2 ൅ 3ܪܰ 3NH

The molar flow rate of oxygen entering the combustor is summed and multiplied by a factor of 

1.25 in order to combust with 25% excess air as shown in equation 29 below. 

ሶܯ

ைଶ,௜௡ ൌ 1.25 · ሺܯሶ஼ு஺ோ,௜௡ ൅ 0.5ܯሶ஼ை,௜௡+ 0.5ܯሶுଶ,௜௡ ൅ 2ܯሶ஼ுସ,௜௡ ൅ 3.5ܯሶ஼ଶு଺,௜௡ ൅

ሶܯ3

஼ଶுସ,௜௡ ൅ 2.5ܯሶ஼ଶுଶ,௜௡ ൅ 5ܯሶ஼ଷு଼,௜௡ ൅ 6.5ܯሶ஼ସுଵ଴,௜௡ ൅ 8ܯሶ஼ହுଵଶ,௜௡ ൅

ሶܯ9.5

஼଺ுଵସ,௜௡ ൅ 11ܯሶ஼଻ுଵ଺,௜௡ ൅ 12.5ܯሶ஼଼ுଵ଼,௜௡ ൅ 14ܯሶ஼ଽுଶ଴,௜௡ ൅

ሶܯ16.5

்஺ோ,௜௡+ 1.5ܯሶுଶௌ,௜௡ ൅ 1.75ܯሶேுଷ,௜௡ሻ

(eqn. 29)

O2TURB 

This block calculates the molar flow rate of air (oxygen and nitrogen) required to combust 

syngas obtained from FT synthesis and the fuel gas obtained from Area 500 in the gas turbine of 

Area 600. A value of 25% excess air is assumed. The calculations are similar to the methodology 

in O2COMB. 

OXYSET 

This block sets the entering oxygen at 0.35 lb oxygen per lb dry biomass into the gasifier. 

݉ሶைଶ,௚௔௦ ൌ 0.35/100 · ݉ሶ ௕௜௢௠௔௦௦ (eqn. 30)

SWGSSTM 
 

78 

This block sets the steam flow into the sour water-gas-shift reactor to be at a ratio of 3:1 water to 

carbon monoxide. This ratio ensures enough water-gas-shift activity occurs within the reactor. 

݉ሶௌ்ெ,௔ௗௗ௜௧௜௢௡ ൌ 3.0 · ݉ሶ ஼ை െ ݉ሶ ுଶை (eqn. 31)
 

79 

Low-Temperature Scenario 

AMINE 

This block calculates the molar flow of monoethanolamine (MEA) needed for the required acid 

gas removal (CO2 and H2S) arriving from syngas quench and FT unconverted syngas recycle 

stream. The MEA is able to capture 0.35 moles acid gas per mole MEA. Additionally, the MEA 

is diluted as explained in DILUTH2O.

Molar MEA flow (in kmol/h) is calculated by 

ሶܯ

ொ஺ ൌ ሺܯሶ஼ைଶ,௦௬௡ ൅ ܯሶ஼ைଶ,௥௘௖ ൅ ܯሶுଶௌ,௦௬௡ሻ/0.35 (eqn. 32)

where ܯሶ

஼ைଶ,௦௬௡ is molar flow of CO2 from the syngas after the syngas quench, ܯሶ஼ைଶ,௥௘௖ is the 

molar flow of CO2 from the unconverted syngas recycle after the FT synthesis, and ܯሶ

ுଶௌ,௦௬௡ is 

the molar flow rate of H2S from the syngas quench. 

Because the MEA solution in the amine absorption unit is to be 20 wt% concentrated with water, 

the flow of water must be calculated. 

Mole flow of water is calculated as 

ሶܯ

ுଶை ൌ ܯሶ

/0.20஺ொܹܯ כ ஺ொ

ܯܹுଶை (eqn. 33)

BIOELEM 

Same as for the HT scenario. 

DILUTH2O 

This block sets the MEA solution to be 20% concentrated with water. 

FTDISTR 

Same as for the high-temperature scenario. 

GASYIELD 

The following model describes how the fluidized-bed gasifier keeps an elemental mass balance. 

Experiments performed at Iowa State University provide the initial gasifier product distribution, 
 

80 

and the model adjusts the yields of those experiments in order to balance carbon, hydrogen, 

sulfur, nitrogen, oxygen, and ash. 

The approach taken to balance each element across the gasifier is by “floating” a component of 

each element. The “floating” component for carbon is the char. All sulfur and nitrogen not found 

in the char is assumed to form hydrogen sulfide and ammonia, respectively. Therefore, sulfur 

and nitrogen are balanced. Next, elemental hydrogen is adjusted in the model by either 

converting diatomic hydrogen to steam or decomposing steam to diatomic hydrogen. Oxygen 

balance is more complex. Because gasification operates at fuel-rich conditions, diatomic oxygen 

should not be present in the syngas leaving the gasifier. Therefore, diatomic oxygen cannot be 

the “floating” component. Instead, oxygen is balanced by adjusting the carbon monoxide or 

carbon dioxide in the exiting syngas. Because there is one oxygen difference between those two 

components, the oxygen can be adjusted to help close the balance. 

Carbon balance follows the flow chart shown in Figure C-4. If there is less gaseous carbon out 

than total carbon in, then the difference is made up of char carbon, CCARB. Char is assumed to 

be composed of 68% carbon with the rest as H, O, N, and S. Ash is considered apart from the 

char and is considered inert in the model. Because the char is now fixed, the only pathway for 

sulfur and nitrogen to take is to form hydrogen sulfide and ammonia. Therefore, the sulfur and 

nitrogen are balanced. 

Figure C-4. Decision diagram for carbon balance 

Next, as shown in Figure C-5, hydrogen is balanced. Knowing hydrogen in the char and in 

gaseous products, the hydrogen required (HREQD) is calculated as the sum of those two 

components. If the hydrogen required is less than hydrogen available (HAVAIL), made up of 

hydrogen in steam, biomass moisture, and in the biomass itself (THYD), then there is enough 

hydrogen available to balance. To balance hydrogen, the product yield swings toward either 

steam or diatomic hydrogen. 
 

81 

Figure C-5. Decision diagram for hydrogen balance 

The only element left to balance is oxygen, which is accomplished by forcing creation of carbon 

monoxide or creation of carbon dioxide as shown in Figure C-6. The required oxygen (OREQD), 

made up of oxygen in char and oxygen in syngas, is checked against the available oxygen found 

in the entering oxygen, steam, and biomass. If there is more oxygen available than required, then 

the option is to move the excess oxygen to CO2 by decreasing CO. If there is still oxygen present 

when CO is decreased to zero, then the yields need to be adjusted because excess oxygen is still 

present. If there is an oxygen deficit (OREQD > OAVAIL), then CO is increased and CO2 is 

decreased. After that, if there is still an oxygen deficit, then insufficient oxygen is present and 

yields need to be adjusted. When all these steps are completed and no errors are generated, there 

is an elemental mass balance across the gasifier. 
 

82 

Figure C-6. Decision diagram for oxygen balance 

GRIND 

This block calculates the power requirement (in kW) for grinding the biomass from the chop size 

of 12 mm to the final size of 6 mm. This power requirement data is found in Mani et al. and is 

for 12% exiting moisture. The correlation has been changed from a polynomial regression 

(which Mani et al. used) to a power regression because the power regression fit the data better. 

ܵ௖௨௧ is in millimeters. 

ܲ௚௥௜௡ௗ ൌ ൫28.76 ڄ ܵ௖௨௧

ି଴.଼ଵ൯ڄ݉ሶ ௕௜௢௠௔௦௦ (eqn. 34)

HUMIDITY 

This block sets humidity of the air entering the air separation unit. 

HV-101, HV-203, HV-445 

This block calculates the lower and higher heating values of the following streams: biomass, 

syngas, and FT products. 

MOISTURE 

This block is the same as found in the HT scenario. 
 

83 

O2COMB 

This block is the same as found in the HT scenario. 

O2TURB 

This block is the same as found in the HT scenario. 

OXYSET 

This block sets the entering oxidizing agents, oxygen and steam, into the gasifier. A linear 

correlation with temperature, ܶ௚௔௦ (in Fahrenheit), adapted from Bain for oxygen, is used 

because as oxygen increases in the gasifier, the temperature increases. Mass flow of oxygen, 

݉ሶைଶ,௚௔௦ , is in percentage of dry feedstock. 

݉ሶைଶ,௚௔௦ ൌ ሺെ11.567 ൅ 0.02375 · ܶ௚௔௦ሻ/100 · ݉ሶ ௕௜௢௠௔௦௦ (eqn. 35)

The steam feed rate is set at 0.66 lb steam per lb oxygen. 

݉ሶ ௦௧௘௔௠,௚௔௦ ൌ 0.66 · ݉ሶைଶ,௚௔௦ (eqn. 36)

Because 95% purity oxygen is produced in the air separation unit, argon mass flow is set at 5% 

of molar oxygen flow. 

݉ሶ ௔௥௚௢௡ ൌ 0.05 · ൬

݉ሶைଶ,௚௔௦

ைଶܹܯ

(37 .eqn஺௥ (ܹܯ כ ൰
 

84 

Aspen Plus Model Design Specifications 

High-Temperature Scenario 

DS-1 

The exiting temperature of air in the heat exchanger used to pre-cool the air entering the 

cryogenic distillation column is varied until a net duty of zero is observed. 

FSSPL02 

This design specification varies the fraction of unconverted syngas that is piped to Area 200 for 

the combustion of syngas. The syngas, in turn, provides the heat required to dry the biomass. 

H2SPLIT 

This design spec calculates the required hydrogen that needs to be reserved by the PSA unit for 

use in Area 500: Hydrocracking. A typical yield from hydrocracking is shown in Table C-4. 

Because the FT products are hydrogen deficient relative to the final blend, makeup hydrogen is 

required. The syngas purge amount going to the pressure swing adsorption (PSA) unit is varied 

so that the calculated delivered hydrogen matches the required hydrogen to Area 500. Without 

showing the detailed calculations, the basic steps are first calculating the carbon and hydrogen 

content in the FT product stream. The carbon mass flow is the same as that of the final blend 

stream flow. Using the blend fractions in Table C-4, the amount of hydrogen is calculated in the 

final blend and the difference in hydrogen is determined. The difference is multiplied by 1.1 to 

obtain the delivered hydrogen mass flow rate to hydrocracking area. 

Table C-4. Hydroprocessing Product Blend 

Component Mass Fraction 

Fuel Gas (methane) 0.034 

LPG (propane) 0.088 

Gasoline (n-octane) 0.261 

Diesel (n-hexadecane) 0.617 

O2-101, O2-203, O2-445 

These design specifications vary the amount of oxygen inlet to the heating value blocks (HV101, HV-203, HV-445) so as to be stoichiometric in the combustion of the duplicate stream. 

O2-SULF 

This design specification varies the amount of oxygen into the LO-CAT oxidizer unit to fully 

oxidize the H2S into solid sulfur. 

SGSTEMP 
 

85 

The temperature of operation in the sour water-gas-shift reactor is varied until the exiting 

equilibrium molar ratio of H2/CO is just above the optimal FT ratio (2.1). A small amount of 

hydrogen is captured in the PSA unit bringing that ratio down to the optimum for FT synthesis. 

Low-Temperature Scenario 

DS-1 

This design specification is the same as for the HT scenario. 

H2SPLIT 

This design specification is the same as for the HT scenario. 

O2-101, O2-203, O2-445 

These design specifications are the same as in the HT scenario. 

O2-SULF 

This design specification is the same as for the HT scenario. 

STMRECOV 

Heat can be recovered from the combustion of syngas and char. This specification varies the 

steam flow rate (stream 280) to bring the combustion flue gas (stream 252) down to 200°C via 

heat exchanging. 

WGSTEMP 

The temperature of operation in the water-gas-shift reactor is varied until the exiting equilibrium 

molar ratio of H2/CO is just above the optimal FT ratio (2.1). A small amount of hydrogen is 

captured in the PSA unit, bringing that ratio down to the optimum for FT synthesis. 
86 

Detailed Calculations 

Aspen Plus Model Calculations and Notes 

 

Outline Defining Units 

Plant Input 

 

Plant Output 

Carbon Efficiency to Fuels 

Energy Content 

FT Reaction Conversion Solver 

Equipment Sizing 

Dryer 

Lockhoppers 

Slag/Char Collection 

PSA Unit 

Fuel Storage 

LT Gasifier Cost 

FT Reactor Cost 

Acid Gas Removal Area Cost 

A500 Hydroprocessing Area Cost 

Reactors and Catalysts 

Natural Gas Utility Usage 

MJ 106:= J MMcf 106ft3:=

kPa 103:= ⋅Pa Cp 100poise :=

ρ water 1000 kg

m3

:= MWH2O 18.02 gm

mol

:=

kmol 1000mo := l MMBTU 106

:= BTU

lbmol kmol

2.2

:=

bbl 42gal :=

ρ gas 737.22 kg

m3

:= 100 kg

m3

6.243

lb

ft3 =

ρ diesel 840 kg

m3

:= therm 100000 BTU :=

MMgal 106:= gal dekatherm 10therm :=

kJ 1000 := J Pref ≡ 1atm

T ref ≡ 298 K bpsd 42gal

day

:=

PJ 1015:= J GJ 10 9

:= J

HHVstover 7.588 103×

BTU

lb =

HHVstover 17.65

MJ

kg

:=
87 

Plant Input 

 

Biomass 

Elemental Composition 

Carbon 

Oxygen 

Hydrogen 

Sulfur 

Nitrogen 

Ash 

Elemental Mass Flow 

mdot_biomass 2000

tonne

day

:= Availability 310 day := Load 7446 hr :=

Frac C_biomass ≡ 0.4728 MWC 12.01 gm

mol

:=

Frac O_biomass ≡ 0.4063 MWO 16. gm

mol

:=

Frac H_biomass ≡ 0.0506 MWH 1.01 gm

mol

:=

Frac S_biomass ≡ 0.0022 MWS 32.07 gm

mol

:=

Frac N_biomass ≡ 0.008 MWN 14.01 gm

mol

:=

Frac A_biomass ≡ 0.0600

mdot_C_in mdot_biomass FracC_biomass := ⋅ mdot_C_in 945.6

tonne

day =

mdot_O_in mdot_biomass FracO_biomass := ⋅ mdot_O_in 812.6

tonne

day =

mdot_H_in mdot_biomass FracH_biomass := ⋅ mdot_H_in 101.2

tonne

day =

mdot_S_in mdot_biomass FracS_biomass := ⋅ mdot_S_in 4.4

tonne

day =

mdot_N_in mdot_biomass FracN_biomass := ⋅ mdot_N_in 16

tonne

day =

mdot_A_in mdot_biomass FracA_biomass := ⋅ mdot_A_in 120

tonne

day =
88 

 

Elemental Mole Flow 

Biomass Moisture 

Source: Kaliyan and Morey, 2005 for 0.66-0.8 mm sized particles 

ndot_C_in

mdot_C_in

MWC

:= ndot_C_in 911.278

mol

s =

ndot_O_in

mdot_O_in

MWO

:= ndot_O_in 587.818

mol

s =

ndot_H_in

mdot_H_in

MWH

:= n dot_H_in 1160

mol

s =

ndot_S_in

mdot_S_in

MWS

:= n dot_S_in 1.588

mol

s =

ndot_N_in

mdot_N_in

MWN

:= n dot_N_in 13.218

mol

s =

moist in := 0.25 moist dried := 0.10

mdot_moist_in

moistin mdot_biomass ⋅

1 moistin − := mdot_moist_in 666.667

tonne

day =

mdot_moist_dried

moistdried mdot_biomass ⋅

1 moistdried − := mdot_moist_dried 222.222

tonne

day =

ρ bulk_stover 100 kg

m3

:=
89 

 

HT Gasifier Steam/Oxygen Addition Source: Probstein and Hicks, 2006

Stoichiometric/thermo-neutral requirement for synthesis gas according to following equation: 

1.34C + 0.34 O2 + H2O --> 0.34CO2 + CO + H2

Oxygen to Carbon: 0.25 

Steam to Carbon: 0.75 

Steam addition ratio is then three times that of oxygen minus the moisture in the biomass 

mdot_O2_in 0.35 mdot_biomass := ⋅ mdot_O2_in 700

tonne

day =

ndot_O2_in

mdot_O2_in

2 MWO ⋅ := ndot_O2_in 253.183

mol

s =

RatioO2_to_C

ndot_O2_in

ndot_C_in

:= RatioO2_to_C = 0.278

ndot_H2O_in 3 RatioO2_to_C ⋅ ndot_C_in ⋅

mdot_moist_dried

MWH2O

:= −

mdot_H2O_in ndot_H2O_in MW H2O := ⋅

ndot_H2O_in 616.817

mol

s =

mdot_H2O_in 960

tonne

day =
90 

Plant Output 

 

HT Fuel Production 

 

 

 

 

LT Fuel Production 

 

 

 

 

mdot_gasHT 112.78

tonne

day

:= mdot_dieselHT 266.5

tonne

day

:=

vdot_gasHT

mdot_gasHT

ρgas

:= vdot_dieselHT

mdot_dieselHT

ρdiesel

:=

vdot_gasHT 40413 gal

day = vdot_dieselHT 83812 galday =

vdot_gasHT 962

bbl

day = vdot_dieselHT 1996

bbl

day =

vdot_gasoline_per_year vdot_gasHT := ⋅Load

vdot_diesel_per_year vdot_dieselHT := ⋅Load

mdot_gasLT 87.12

tonne

day

:= mdot_dieselLT 205.86

tonne

day

:=

vdot_gasLT

mdot_gasLT

ρgas

:= vdot_dieselLT

mdot_dieselLT

ρdiesel

:=

vdot_gasLT 31218 gal

day = vdot_dieselLT 64741 galday =

vdot_gasLT 743

bbl

day = vdot_dieselLT 1541

bbl

day =

vdot_gasoline_per_yearLT vdot_gasLT := ⋅Load vdot_gasoline_per_yearLT = 9.685 MMgal

vdot_diesel_per_yearLT vdot_dieselLT := ⋅Load vdot_diesel_per_yearLT = 20.086 MMgal

vdot_gasoline_per_year = 12.538 MMgal

vdot_diesel_per_year = 26.003 MMgal
91 

Carbon Efficiency to Fuels 

 

HT Scenario 

Gasoline Carbon 

Diesel Carbon 

mdot_gasHT 112.78

tonne

day =

mdot_C_gasHT FracC_gasoline mdot_gasHT := ⋅ mdot_C_gasHT 94.835

tonne

day =

mdot_dieselHT 266.5

tonne

day =

FracC_diesel

16 12.01 ⋅

16 12.01 ⋅ 34 1.01 ⋅+

:= FracC_diesel = 0.848

mdot_C_dieselHT FracC_diesel mdot_dieselHT := ⋅ mdot_C_dieselHT 226.096

tonne

day =

mdot_C_outHT mdot_C_gasHT mdot_C_dieselHT := +

mdot_C_outHT 320.931

tonne

day = C_effHT

mdot_C_outHT

mdot_C_in

:=

FracC_gasoline

8 12.01 ⋅

8 12.01 ⋅ 18 1.01 ⋅+

:= FracC_gasoline = 0.841

C_effHT 0.339 =
92 

 

LT Scenario 

Gasoline Carbon 

Diesel Carbon 

mdot_gasLT 87.12

tonne

day =

mdot_C_gasLT FracC_gasoline mdot_gasLT := ⋅ mdot_C_gasLT 73.258

tonne

day =

mdot_dieselLT 205.86

tonne

day =

mdot_C_dieselLT FracC_diesel mdot_dieselLT := ⋅ mdot_C_dieselLT 174.649

tonne

day =

mdot_C_outLT mdot_C_gasLT mdot_C_dieselLT := +

mdot_C_outLT 247.908

tonne

day = C_effLT

mdot_C_outLT

mdot_C_in

:= C_effLT 0.262 =
93 

Egasifierloss_LT = 96.796 MW

Egasifierloss_HT = 47.111 MW

Energy Content 

This section acquires the energy content (on a LHV basis) from the Aspen Plus data and converts to 

megawatts for use in developing an energy balance 

Biomass 

Fuel 

Char/Tar 

Raw Syngas 

Energy Loss across the Gasifier 

Energy lost across the gasifier is calculated as difference in energy in the biomass and energy 

in the raw syngas and char (only in LT scenario) 

Ebiomass 1400313

MJ

hr

:= Ebiomass = 388.976 MW

EfuelHT 695598

MJ

hr

:= EfuelHT = 193.222 MW

EfuelLT 539292

MJ

hr

:= EfuelLT = 149.803 MW

Echar_LT 87792

MJ

hr

:= Echar_LT = 24.387 MW

Etar_LT 16980

MJ

hr

:= Etar_LT = 4.717 MW

Erawsyngas_HT 1230712

MJ

hr

:= Erawsyngas_HT = 341.864 MW

Erawsyngas_LT 964054

MJ

hr

:= Erawsyngas_LT = 267.793 MW

Egasifierloss_HT Ebiomass Erawsyngas_HT := −

Egasifierloss_LT Ebiomass Erawsyngas_LT − Echar_LT := −
94 

 

Unconverted Syngas used in A600 Power Generation 

Fuel Gas from A500 used in A600 Power Generation 

Fischer-Tropsch Product 

Electricity Generated 

Net Electricity (exported) 

Power Generation Loss 

The loss is the difference between electric generation out and the gas energy in 

EsyngasA600_HT 129332

MJ

hr

:= EsyngasA600_HT = 35.926 MW

EsyngasA600_LT 109708

MJ

hr

:= EsyngasA600_LT = 30.474 MW

Efuelgas_HT 104114

MJ

hr

:= Efuelgas_HT = 28.921 MW

Efuelgas_LT 80718

MJ

hr

:= Efuelgas_LT = 22.422 MW

EFTliquids_HT 782894

MJ

hr

:= EFTliquids_HT = 217.471 MW

EFTliquids_LT 606801

MJ

hr

:= EFTliquids_LT = 168.556 MW

EelecgenOUT_HT := 48.55 MW EelecgenOUT_LT := 40.73 MW

Eelecnet_HT := 13.8 MW Eelecnet_LT := 16.3 MW

EA600losses_HT EsyngasA600_HT Efuelgas_HT + EelecgenOUT_HT := −

EA600losses_HT = 16.296 MW

EA600losses_LT EsyngasA600_LT Efuelgas_LT + EelecgenOUT_LT := −

EA600losses_LT = 12.166 MW
95 

 

Loss across FT Reactor 

Unconverted Syngas used for Biomass Drying 

Only in HT scenario 

EFTreactorlosses_HT 226737

MJ

hr

:= EFTreactorlosses_HT = 62.983 MW

EFTreactorlosses_LT 175128

MJ

hr

:= EFTreactorlosses_LT = 48.647 MW

Ebiomass_drying_HT 24663

MJ

hr

:= Ebiomass_drying_HT = 6.851 MW
96 

Fischer-Tropsch Reaction Conversion Solver 

 

This section solves for the reaction fractional conversion for each reaction in the FischerTropsch reactor. A set of equations is developed and solved. The resulting ε values (ε1 - ε30) 

are used directly in the Aspen Plus conversion reactor block. The reactions in the reactor block 

are defined as molar extent. 

Depending on the alpha chain growth probability, the reactor forms different product 

composition. 

Step 1: choose the expected alpha chain growth value 

Step 2: using the αFT chain growth, the mole fraction of each hydrocarbon in the FT product is 

calculated. 

 

 

 

 

 

 

 

 

 

All hydrocarbons greater than C20 make up the balance and are modeled using C30. 

α FT := 0.9

M1 αFT

1 1− 1 αFT := ⋅( ) − M1 0.1 = M11 αFT11 1− 1 αFT := ⋅( ) −

M2 αFT

2 1− 1 αFT := ⋅( ) − M2 0.09 = M12 αFT12 1− 1 αFT := ⋅( ) −

M3 αFT

3 1− 1 αFT := ⋅( ) − M3 0.081 = M13 αFT13 1− 1 αFT := ⋅( ) −

M4 αFT

4 1− 1 αFT := ⋅( ) − M4 0.073 = M14 αFT14 1− 1 αFT := ⋅( ) −

M5 αFT

5 1− 1 αFT := ⋅( ) − M5 0.066 = M15 αFT15 1− 1 αFT := ⋅( ) −

M6 αFT

6 1− 1 αFT := ⋅( ) − M6 0.059 = M16 αFT16 1− 1 αFT := ⋅( ) −

M7 αFT

7 1− 1 αFT := ⋅( ) − M7 0.053 = M17 αFT17 1− 1 αFT := ⋅( ) −

M8 αFT

8 1− 1 αFT := ⋅( ) − M8 0.048 = M18 αFT18 1− 1 αFT := ⋅( ) −

M9 αFT

9 1− 1 αFT := ⋅( ) − M9 0.043 = M19 αFT19 1− 1 αFT := ⋅( ) −

M10 αFT

10 1− 1 αFT := ⋅( ) − M10 0.039 = M20 αFT20 1− 1 αFT := ⋅( ) −

M30 1 M1 M2 + + M3 + M4 + M5 + M6 + M7 + M8 + M9 + M10

+ M11 M12 + + M13 + M14 + M15 + M16 + M17 + M18 + M19 + M20

 ...

 

:= −

M30 0.122 =

M13 0.028 =

M14 0.025 =

M11 0.035 =

M12 0.031 =

M15 0.023 =

M16 0.021 =

M17 0.019 =

M18 0.017 =

M19 0.015 =

M20 0.014 =
97 

Step 3: Set up a series of equations to solve along with guess values (required for Mathcad) 

For a nominal 1000 moles of CO input, the expected CO output is 600 moles because 40% is 

converted. 

<----------------- 40% conversion of CO 

 

 

 

 

<----- This value to be varied until COconv is equal to desired. 

A nominal 400 moles of CO are converted in the FT reactor. The sum of the exiting amount of 

moles in the FT product distribution will not be 400, because moles are not conserved. Mass is 

conserved, however. Therefore, the variable "D" represents a factor that adjusts all the 

conversions (ε1, ε2, etc.). 

The resulting value of D is 0.1 meaning that 40 moles of FT products exit the reactor. 

 

 

 

 

Known CO out := 600 CO in := 1000

Guess ε1 20 := ε2 20 := ε3 20 := ε4 20 := ε5 20 :=

ε6 20 := ε7 20 := ε8 20 := ε9 20 := ε10 20 :=

ε11 20 := ε12 20 := ε13 20 := ε14 20 := ε15 20 :=

ε16 20 := ε17 20 := ε18 20 := ε19 20 := ε20 20 :=

ε30 20 := D 0.1 :=

Given

D ε1

1

2

+ ε2

1

3

+ ε3

1

4

+ ε4

1

5

+ ε5

1

6

+ ε6

1

7

+ ε7

1

8

+ ε8

1

9

+ ε9

1

10

+ ε10

1

11

+ ε11

1

12

+ ε12

1

13

ε13

1

14

+ ε14

1

15

+ ε15

1

16

+ ε16

1

17

+ ε17

1

18

+ ε18

1

19

+ ε19

1

20

+ ε20

1

30

+ + ε30

... 



















M1 ε1

D

M2

1

2

ε2

D

M3

1

3

ε3

D

M4

1

4

ε4

D

M5

1

5

ε5

D

M6

1

6

ε6

D

M7

1

7

ε7

D

M8

1

8

ε8

D

M9

1

9

ε9

D

M10

1

10

ε10

D

M11

1

11

ε11

D

M12

1

12

ε12

D

M13

1

13

ε13

D

M14

1

14

ε14

D

M15

1

15

ε15

D

M16

1

16

ε16

D

M17

1

17

ε17

D

M18

1

18

ε18

D

M19

1

19

ε19

D

M20

1

20

ε20

D

M30

1

30

ε30

D

Solve Find := ( ) ε1, ε ε2, ε 3, ε 4, ε 5, ε 6, ε 7, ε 8, ε 9, ε 10, ε 11, ε 12, ε 13, ε 14, ε 15, ε 16, ε 17, ε 18, ε 19, ε 20, 30

Solve

0

0

1

0.01

0.018

=
98 

 

Step 4: The guess value of D is varied until the sum of all reaction conversions (ε1, ε2, etc.) 

sum to 1.0 as seen below. This means that all 400 moles of CO are converted as expected. 

Step 5: Each value for ε is imported into Aspen Plus 

e1 Solve0 := e1 0.01 = e11 Solve 10 := e11 0.03835 = D 0.1 =

e2 Solve1 := e2 0.018 = e12 Solve 11 := e12 0.03766 =

e3 Solve2 := e3 0.0243 = e13 Solve 12 := e13 0.03672 =

e4 Solve3 := e4 0.02916 = e14 Solve 13 := e14 0.03559 =

e5 Solve4 := e5 0.0328 = e15 Solve 14 := e15 0.03432 =

e6 Solve5 := e6 0.03543 = e16 Solve 15 := e16 0.03294 =

e7 Solve6 := e7 0.0372 = e17 Solve 16 := e17 0.0315 =

e8 Solve7 := e8 0.03826 = e18 Solve 17 := e18 0.03002 =

e9 Solve8 := e9 0.03874 = e19 Solve 18 := e19 0.02852 =

e10 Solve 9 := e10 0.03874 = e20 Solve 19 := e20 0.02702 =

e30 Solve 20 := e30 0.36473 =

COconv e1 e2 + + e3 + e4 + e5 + e6 + e7 + e8 + e9 + e10 + e11 + e12 + e13 + e14 + e15 + e16

+ e17 e18 + + e19 + e20 + e30

:= ...

COconv = 1
99 

Equipment Sizing 

 

Rotary Dryer Source: Process Engineering Economics by James Couper, 2003 

Typical rpm of rotary dryers 

Typical product of equals 15-25. Assume value of 25 for larger end 

Typical residence times are 5-90 minutes and holdup of solids is 7%-8%. Assume 5 minutes 

and 8%. 

Typical exit gas temperature is 10°-20°C above the entering solids. 

Feed rate into plant is 2000 ton/day with bulk density of stover equal to 100kg/m^3. Water 

density is accounted for as well. 

 

Volume of solids in dryer 

Volume of solids and steam 

Length of theoretical dryer 

Surface area of theoretical dryer 

Max surface area as reported by Aspen Icarus is 185 m

2

; therefore approximately 10 dryers are 

required. 

Feed throughput in each dryer (used for Icarus input) 

rpm dryer := 4

rpm diameter feet ⋅ ( )

Ddryer

25ft

rpmdryer

:= Ddryer = 6.25 ft

t

res := 5min holdup 0.08 :=

mdot_feed 2000 tonne

day

:= mdot_moist_in 666.667

tonne

day =

ρbulk_stover 100 kg

m3 = ρwater 1000 kgm3 =

Vsolids

mdot_feed

ρ bulk_stover

mdot_moist_in

ρ water

+













t

res := ⋅

Vdryer_total

Vsolids

holdup :=

lengthdryer

Vdryer_total

Ddryer

2 π

4

⋅

:=

Asurf_dryer length dryer π⋅ Ddryer := ⋅

mdot_feed mdot_moist_in +

10

24495.8

lb

hr =

Vsolids 71.759m3 =

Vdryer_total 896.991m3 =

length dryer = 314.708 m

Asurf_dryer 1883.4m2 =
100 

 

Lockhopper System 

Source: CE IGCC Repowering Project Bins and Lockhoppers, Combustion Eng. 1993 

Note: this report's feedstock is coal 

Assumptions from report 

-A receiving bin is situated before the lockhopper with a 40 minute residence time 

-Design pressure is for 50 psia. 

-Cycle time for lockhopper system is designed for 10 minutes resulting in approximately 50,000 

cycles per year 

-Storage volume for lockhopper and feed bin is assumed to be 10 minutes 

-Approximate lockhopper and feed bin vessel thickness is 1.5 inches and design pressure is for 

450 psia 

-Volume is theoretical + 33% 

Residence Time 

biomass receiving bin 

biomass lockhopper 

biomass feed bin 

Density of feed 

HT Scenario Lockhopper System (1 train) 

Volume of biomass receiving bin 

Volume of biomass lockhopper 

Volume of biomass feed bin 

t

res_rbin := 40min ε void := 25%

t

res_lock := 10min

t

res_fbin := 10min

mdot_feed_lock mdot_feed mdot_moist_dried := +

mdot_feed_lock 2222 tonne

day =

ρ stover_10%moist

ρ bulk_stover ⋅ ρ 2000 water⋅+ 222

2222

:=

ρ stover_10%moist 189.919 kg

m3 =

Vr_bin

t

res_rbin mdot_feed_lock ⋅

ρstover_10%moist

1

1 εvoid − := ⋅

Vlock

t

res_lock mdot_feed_lock ⋅

ρstover_10%moist

1

1 εvoid − := ⋅

Vf_bin

t

res_fbin mdot_feed_lock ⋅

ρstover_10%moist

1

1 εvoid − := ⋅

Vr_bin 433m3 =

Vlock 108m3 =

Vf_bin 108m3 =
101 

 

LT Scenario Lockhopper System (7 trains) 

Volume of biomass receiving bin 

Volume of biomass lockhopper 

Volume of biomass feed bin 

Source: Combustion Engineering 1993 

Vr_binLT

Vr_bin

7 :=

VlockLT

Vlock

7

:=

Vf_binLT

Vf_bin

7

:=

Vr_binLT 61.909m3 =

VlockLT 15.477m3 =

Vf_binLT 15.477m3 =
102 

Lockhopper Power Consumption 

Source: Techno-Economic Analysis of Hydrogen Production by Gasification of Biomass by 

Lau et al. [2002] 

Specific power of lockhopper, kW/tonne/day 

Biomass inlet to gasifier 

Fly Ash Collection Storage Tank (assume 7 days storage) 

(assumed) 

 

Slag Separation Drum (5 minute residence time, 20% volume) 

Slag Collection Storage Tank (7 days storage) 

Note: the resulting volumes are used to assist in costing 

using Aspen Icarus 

Char Collection Storage Bin (1-day residence time, 80% volume) 

assume 20% voidage 

SPlock 0.082 kW

tonne

day

:=

mdot_gasifier mdot_biomass mdot_moist_dried := +

Power lock SPlock mdot_gasifier := ⋅ Power lock = 182.222 kW

ρ ash 700 kg

m3

:= mdot_ash 5.88

tonne

day

:=

Vtank

mdot_ash

ρash

:= ⋅7day Vtank 58.8m3 = Vtank 2.077 103

× ft3 =

ρ slag 2700 kg

m3

:= mdot_slag 114

tonne

day

:= εvoid_slag := 0.8

Vdrum

mdot_slag

ρslag

⋅5min 1

1 εvoid_slag − := ⋅ Vdrum 0.733m3 =

Vchardrum

mdot_char

ρchar

⋅1day 1

1 εvoid_char − := ⋅

Vslag_storage 295.6m3 =

Vslag_storage

mdot_slag

ρslag

:= ⋅7day

εvoid_char ρ char 2700 := 0.2 kg

m3

:= mdot_char 214

tonne

day

:= Vchardrum 99.074m3 =
103 

 

Pressure Swing Adsorption Unit Sizing 

The adsorption unit is 1/3 molsieve and 2/3 activated carbon 

(a) (b) (b) 

Molsieve 13X 

Determine dry volumetric flow rate of the syngas stream at atmospheric pressure and 25 deg C 

Mole fraction of components that are adsorbed 

 

Actual flow rate of components adsorbed 

Adsorbent capacity 

(d) SCF/lb corrected for P and T to actual 

cm3

/gm; PSA occurs at ambient 

temperature 

Mass of molsieve required 

References in parentheses are 

given at the end of this section. Pi 3.1415 := nm 10− 9

:= ⋅m

BulkDens 43 lb

ft3

:= ⋅ SA 1320 m2

gm

:= ⋅ PoreVol 0.51 cm

3

gm

:= ⋅

VolFlowRate 167 1 ( ) − kmol

hr ⋅ ⋅22.414 m3

kmol

⋅ 14.696 psi ⋅

400 psi ⋅ ⋅ ( )K 273.15 25 + ⋅273.15 K⋅ := ⋅

VolFlowRate 149.211m3

hr =

CO 23 := CO2 1 := CH4 1 :=

FlowRateAds VolFlowRate CO CO2 + + CH4

100

:= ⋅

FlowRateAds 37.303 m3

hr =

AdsCap 0.34 ft3

lb

⋅ 14.696 psi ⋅

400 psi ⋅ ⋅ ( )K 273.15 25 + ⋅273.15 K⋅ := ⋅

AdsCap 0.851 cm

3

gm =

CycleTime 5 mi := ⋅ n

MolSieveMass FlowRateAds CycleTime ⋅

AdsCap

:= MolSieveMass 3.652 103 = × kg
104 

 

Determine volume and length of molsieve bed and activated carbon bed 

(assumed) 

(Just molsieve bed) 

(bed is 1/3 molsieve, 2/3 activated carbon) 

(a) http://www.sigmaaldrich.com/Brands/Aldrich/Tech_Bulletins/AL_143/Molecular_Sieves.html 

(b) US Pat 6117810 

(d) WO/1998/058726 BULK SEPARATION OF CARBON DIOXIDE FROM METHANE 

USING NATURAL CLINOPTILOLITE --extrapolate to partial pressure of 

CO2+CH4+N2+CO=32.6%*400 psi 

BedVolume MolSieveMass

BulkDens

:= BedVolume 5.302 m3 =

Diam 4 f := ⋅ t Diam 1.219 m =

Length BedVolume

Pi Diam2 ⋅

:= Length 3.725f = t

RxtrLength 3 Length := ⋅

RxtrLength 11.175f = t RxtrLength 3.406 m =

RxtrVolume RxtrLength Diam2:= ⋅ ⋅0.25⋅π RxtrVolume 3.977m3 =

×4 
105 

 

HT Scenario Fuel Storage 

Gasoline Storage Tank (30 days storage) 

Diesel Storage Tank (30 days storage) 

Note: the resulting volumes are used to assist in costing using Aspen Icarus software 

mdot_gasHT 112.78

tonne

day =

vdot_gasHT 4.041 104

× gal

day =

Vgas_tankHT vdot_gasHT := ⋅30day

Vgas_tankHT 4589m3 =

mdot_dieselHT 266.5

tonne

day =

vdot_dieselHT 8.381 104

× gal

day =

Vdiesel_tankHT vdot_dieselHT := ⋅30day

Vdiesel_tankHT 9518m3 =
106 

 

LT Scenario Fuel Storage 

Gasoline Storage Tank (30 days storage) 

Diesel Storage Tank (30 days storage) 

Note: the resulting volumes are used to assist in costing using Aspen Icarus 

mdot_gasLT 87.12

tonne

day =

vdot_gasLT 3.122 104× gal

day =

Vgas_tankLT vdot_gasLT := ⋅30day

Vgas_tankLT 3545 m3 =

mdot_dieselLT 205.86

tonne

day =

vdot_dieselLT 6.474 104× gal

day =

Vdiesel_tankLT vdot_dieselLT := ⋅30day

Vdiesel_tankLT 7352 m3 =
107 

LT Gasifier Cost 

 

Source: Larson et al. 2005 in 2003$ 

$MM 

Biomass throughput of 300 tpd 

The cost ($MM) of one train at 300 ton per day 

$MM 

Because 2,205 ton/day we need 7 gasifiers but we can apply the multiple train scaling exponent 

$MM 

C0_gasifier 6.41 106 := ⋅ S0_gasifier 41.7

tonne

hr

:= Smax 120

tonne

hr

:=

SgasifierLT 300

ton

day

:= SgasifierLT 11.34

tonne

hr =

CgasifierLT C0_gasifier

SgasifierLT

tonne

hr

1

S0_gasifier

tonne

hr

⋅ 



















f

:= ⋅

CgasifierLT 2.576 106×=

mtrain := 0.9

CgasifierLTtrain CgasifierLT 7

mtrain := ⋅ CgasifierLTtrain 1.484 107

×=

f 0.7 :=
108 

FT Reactor Costing 

 

Source: Larson et al. 2005 in 2003$ 

$MM 

HT Scenario 

 

Installed cost $MM (assume 3.6 install factor consistent with 

Peters et al.) 

LT Scenario 

 

Installed cost $MM (assume 3.6 install factor consistent with 

Peters et al.) 

CFT_base := 10.5 f

FT2 := 0.72 SFT_base 2.52

MMcf

hr

:=

Mdot_FTHT 13829

kmol

hr

:= Vstandard_FTHT Mdot_FTHT ⋅22.4

L

mol

:=

Vstandard_FTHT 10.939

MMcf

hr =

CFTHT_reac CFT_base

Vstandard_FTHT

SFT_base













f

FT2

:= ⋅

CFTHT_reac = 30.217

Mdot_FTLT 11400

kmol

hr

:= Vstandard_FTLT Mdot_FTLT ⋅22.4

L

mol

:=

Vstandard_FTLT 9.018

MMcf

hr =

CFTLT_reac CFT_base

Vstandard_FTLT

SFT_base













f

FT2

:= ⋅

CFTLT_reac = 26.294
109 

Acid-Gas Removal Area Cost 

 

Source: Phillips et al. 2007 in 2005$ 

Calculated by adding the input syngas streams to the absorber column 

HT Scenario 

 

LT Scenario 

 

SAGR_base 332910

lb

hr

:= fAGR := 0.65 CAGR_base := 5446503

SAGR_HT ( ) 2965 2308 +

tonne

day

:= ⋅ SAGR_HT 484374

lb

hr =

CAGR_HT CAGR_base

SAGR_HT

lb

hr

1

SAGR_base

lb

hr

⋅ 



















f

AGR

:= ⋅ CAGR_HT = 6949808

SAGR_LT ( ) 2070 2190 +

tonne

day

:= SAGR_LT 391321

lb

hr =

CAGR_LT CAGR_base

SAGR_LT

lb

hr

1

SAGR_base

lb

hr

⋅ 



















f

AGR

:= ⋅ CAGR_LT = 6049946
110 

A500 Hydroprocessing Area Cost 

 

Source: Robinson et al. 2007 in 2007$ 

Note: "bpsd" is barrels per standard day 

HT Scenario 

(from Aspen Plus model) 

Power required for A500 

LT Scenario 

(from Aspen Plus model) 

Power required for A500 

AreaCost 0 (assumed) 

4000

bpsd

:= S0_HY := 25000 bpsd

C0_HY AreaCost 0 S0_HY := ⋅ C0_HY = 100000000

mdot_FTL_HT 428

tonne

day

:= ρ FTL 750 kg

m3

:=

vdot_FTL_HT

mdot_FTL_HT

ρFTL

:= vdot_FTL_HT 3.589 103 = × bpsd

CHY_HT C0_HY

vdot_FTL_HT

bpsd

1

S0_HY

bpsd

⋅ 



















f

HY

:= ⋅ CHY_HT 2.832 107

×=

Power per_bpsd

15kW hr ⋅

bpsd day ⋅ :=

PowerareaHT Powerper_bpsd vdot_FTL_HT := ⋅ Power areaHT = 2.243 MW

mdot_FTL_LT 330.42

tonne

day

:=

vdot_FTL_LT

mdot_FTL_LT

ρFTL

:= vdot_FTL_LT 2.771 103 = × bpsd

CHY_LT C0_HY

vdot_FTL_LT

bpsd

1

S0_HY

bpsd

⋅ 



















f

HY

:= ⋅ CHY_LT 2.394 107

×=

PowerareaLT Powerper_bpsd vdot_FTL_LT := ⋅ Power areaLT = 1.732MW

fHY := 0.65
111 

Reactors and Catalysts 

 

Fischer-Tropsch Reactor and Cobalt Catalyst 

FT reactor volume 

Using gas hourly space velocity and actual volumetric flow rate, the volume of the reactor is 

determined 

(assumed) 

(from Aspen Plus model) 

 

 

FT catalyst cost 

(assumed) (assumed) 

Replacement cost of cobalt catalyst 

GHSV

ν 0

V

GHSVFT 100 hr− 1:= ⋅

νrate_actHT 6.298m3

s

:=

VFTHT

νrate_actHT

GHSVFT

:= VFTHT 226.728m3 = VFTHT 8.007 103

× ft3 =

νrate_actLT 5.021m3

s

:=

VFTLT

νrate_actLT

GHSVFT

:= VFTLT 180.756m3 = VFTLT 6.383 103

× ft3 =

Cocost

15

lb

:= ρCo 64

lb

ft3

:=

Covol_cost Cocost ρCo := ⋅ Covol_cost 960

1

ft3 =

Cototal_costHT Covol_cost VFTHT := ⋅ Cototal_costHT 7.687 106

×=

Cototal_costLT Covol_cost VFTLT := ⋅ Cototal_costLT 6.128 106

×=
112 

 

Water Gas Shift Reactor and Catalyst 

Sour WGS reactor volume (HT scenario) 

Using gas hourly space velocity and actual volumetric flow rate, the volume of the reactor is 

determined 

(assumed) 

(from Aspen Plus) 

 

WGS reactor volume (LT scenario) 

(from Aspen Plus) 

 

WGS and SWGS Catalyst Cost 

 

Replacement cost of WGS catalyst 

GHSV

ν0

V

GHSVWGS 1000 hr− 1 := ⋅

νrate_actSWGS 2.008 m3

s

:=

VSWGS

νrate_actSWGS

GHSVWGS

:= VSWGS 7.229m3 = VSWGS 255.283ft3 =

νrate_actWGS 1.834 m3

s

:=

VWGS

νrate_actWGS

GHSVWGS

:= VWGS 6.602m3 = VWGS 233.162ft3 =

CatCostWGS

8

lb

:= ρ cat_WGS 56

lb

ft3

:= ρ cat_WGS 897.034 kg

m3 =

CatCost vol_WGS CatCost WGS ρ cat_WGS := ⋅ CatCost vol_WGS 448

1

ft3 =

TotalCatCost SWGS CatCost vol_WGS VSWGS := ⋅ TotalCatCost SWGS = 114367

TotalCatCost WGS CatCost vol_WGS VWGS := ⋅ TotalCatCost WGS = 104456
113 

 

Steam Methane Reformer Reactor and Catalyst (LT Scenario) 

(assumed) 

(from Aspen Plus model) 

 

SMR Catalyst Cost 

 

Replacement cost of SMR catalyst 

GHSVSMR 2600hr− 1 :=

νrate_actSMR 7.082 m3

s

:=

VSMR

νrate_actSMR

GHSVSMR

:= VSMR 9.806m3 = VSMR 346.29ft3 =

CatCost SMR

4.67

lb

:= ρ cat_SMR 64

lb

ft3

:= ρ cat_SMR 1.025 103× kg

m3 =

CatCost vol_SMR CatCost SMR ρ cat_SMR := ⋅ CatCost vol_SMR 298.88

1

ft3 =

TotalCatCost SMR CatCost vol_SMR VSMR := ⋅ TotalCatCost SMR = 103499
114 

Natural Gas Utility Consumption 

 

Annual natural gas requirement at 5% of yearly operating hours 

Natural gas properties 

 

HT Scenario 

(from Aspen Plus model, includes power required for 

gas turbine air compressor) 

(assumed) 

Annual natural gas requirement 

Average flow rate of natural gas 

Price (Source: Energy Information Administration) 

Cost ng

6.4

1000 ft3

:=

HHVng 54

MJ

kg

:= MWng 16.04 gm

mol

:= ρ ng 22.4

L

mol

:=

Cost ng

ρ ng

MWng

⋅ 286.335

1

ton =

Prequired_plantHT := 32.813MW

Effng_to_power := 0.35

mdot_ngHT

Prequired_plantHT

Effng_to_power

HHVng

:= mdot_ngHT 1.378 104

×

lb

hr =

M ngHT mdot_ngHT := ⋅Availability ⋅0.05 M ngHT = 2563 ton

mdot_ng_5%HT

MngHT

8760hr

:= mdot_ng_5%HT 585.14

lb

hr =
115 

LT Scenario 

(from Aspen Plus model, includes power required for 

gas turbine air compressor) 

Annual natural gas requirement 

Average flow rate of natural gas 

Prequired_plantLT := 24.3MW

mdot_ngLT

Prequired_plantLT

Effng_to_power

HHVng

:= mdot_ngLT 1.02 104

×

lb

hr =

M ngLT mdot_ngLT := ⋅Availability ⋅0.05 M ngLT = 1898 ton

mdot_ng_5%LT

M ngLT

8760 hr

:= mdot_ng_5%LT 433.331

lb

hr =
116 

Appendix D. Process Flow Diagrams 
117 

High-Temperature Scenario 

Figure D-1. Overall plant area process flow diagram for HT scenario 
118 

Figure D-2. Preprocessing area process flow diagram for HT scenario 
119 

Figure D-3. Gasification area process flow diagram for HT scenario 
120 

Figure D-4. Syngas cleaning area process flow diagram for HT scenario 
121 

Figure D-5. Acid gas removal area process flow diagram for HT scenario 
122 

Figure D-6. Sulfur recovery area process flow diagram for HT scenario 
123 

Figure D-7. Fuel synthesis area process flow diagram for HT scenario 
124 

Figure D-8. Hydroprocessing area process flow diagram for HT scenario 
125 

Figure D-9. Power generation area process flow diagram for HT scenario 
126 

Figure D-10. Air separation unit process flow diagram for HT scenario
127 

Low-Temperature Scenario 

Figure D-11. Overall plant area process flow diagram for LT scenario 
128 

Figure D-12. Preprocessing area process flow diagram for LT scenario 
129 

Figure D-13. Gasification area process flow diagram for LT scenario 
130 

Figure D-14. Syngas cleaning area process flow diagram for LT scenario 
131 

Figure D-15. Acid gas removal area process flow diagram for LT scenario 
132 

Figure D-16. Sulfur recovery process flow diagram for LT scenario 
133 

Figure D-17. Fuel synthesis area process flow diagram for LT scenario 
134 

Figure D-18. Syngas conditioning area process flow diagram for LT scenario 
135 

Figure D-19. Hydroprocessing area process diagram for LT scenario 
136 

Figure D-20. Power generation area process flow diagram for LT scenario 
137 

Figure D-21. Air separation unit process flow diagram for LT scenario
138 

Appendix E. Stream Data 
139 

High-Temperature Scenario 

Table E-1. Overall Plant Stream Data for HT Scenario 

HT Overall Plant

PL00BMAS PL08BMAS PL09CAIR PL17SLAG PL21SGAS PL34SGASPL42CO2 PL43CO2 PL47SGAS PL48SGAS PL49SGASPL50FT PL52FLUE PL65FGAS PL71GASOPL81DIESPL81STMPL81WATPL83SUL PL84STM PL88FLUEPL89AIRPL90HYDPL90OX

Temperature (C) 25 90 25 50 1300 62 250 53 45 45 45 35 220 35 36 37 200 40 50 120 273 32 30 Pressure (bar) 1.01 1.01 1.01 26.62 26.62 22.75 28.00 3.45 23.58 23.58 23.58 22.20 1.03 22.20 1.03 1.03 1.98 24.82 1.93 1.98 1.00 1.01 1.00 28

Vapor Fraction 0.00 0.00 1.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 0.00 0.00 1.00 0.00 0.00 1.00 1.00 1.00 1.00 

Volume Flow** (m3/sec) 0.43 0.14 2.33 0 11.23 3.45 0.07 3.32 0.02 1.64 0.11 0.01 4.33 0.02 0 0.01 50.53 0.02 0 42.34 45.37 29.07 0.63 

Mole Flow** (kmol/hr) 1542 513.97 343.36 0 8192 10089 175.05 1541 71.82 5260 355.46 89.49 393.42 74.55 41.08 48.97 9251 3276 13.46 9251 3595 4177 90.27 95

Mass Flow (tonnes/day) 2667 2222 237.82 114.00 3825 3377 180.00 1585 29.08 2130 143.95 427.14 266.91 52.77 112.62 266.11 4000 1501 7.20 4000 2439 2903 4.37 74 H2O 666.67 222.22 0 0 988.43 45.53 2.50 21.97 0 0 0 0 27.89 0 0 0 0 1348 4.04 0 233.43 0 0  CO 0 0 0 0 1457 1818 0.47 4.15 11.24 823.54 55.65 0 0 0 0 0 0 22.10 0 0 0 0 0  H2 0 0 0 0 122.88 288.34 0 0 1.83 134.08 9.06 0 0 0 0 0 0 0 0 0 0 0 4.37 

 CO2 0 0 0 0 1184 190.01 175.21 1543 2.40 175.74 11.88 0 40.00 0 0 0 0 128.26 0 0 352.24 0 0  O2 0 0 55.97 0 0 0 0 0 0 0 0 0 10.28 0 0 0 0 0 0 0 104.25 672.19 0 70 N2 0 0 181.85 0 17.68 0 1.78 15.63 0 0 0 0 181.85 0 0 0 0 0.26 0 0 1715 2194 0  CH4 0 0 0 0 0.02 63.41 0.01 0.07 0.87 63.47 4.29 0 0 14.94 0 0 0 0 0 0 0 0 0  C2H6 0 0 0 0 0 106.90 0.01 0.10 1.46 107.01 7.23 0 0 0 0 0 0 0 0 0 0 0 0  C2H4 00000000000000000000000 C2H2 00000000000000000000000 C3 0 0 0 0 0 141.45 0.01 0.10 1.93 141.56 9.57 0 0 37.83 0 0 0 0 0 0 0 0 0 

 C4 0 0 0 0 0 167.38 0.02 0.15 2.29 167.55 11.32 0 0 0 0 0 0 0 0 0 0 0 0 

 H2S 0 0 0 0 4.50 0.03 0 0 0 0 0 0 0 0 0 0 0 1.02 0 0 0 0 0  NH3 0 0 0 0 0.11 0 0 0 0 0 0 0 0 0 0 0 0 0.10 0 0 0 0 0  TAR 00000000000000000000000 SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 3.16 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 00000000000000000000000 STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 4000 0 0 4000 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0.08 0 0  NO2 00000000000000000000000 MEA 00000000000000000000000 AIR 00000000000000000000000 WAXES 0 0 0 0 0 0 0 0 0 0 0 170.22 0 0 0 0 0 0 0 0 0 0 0 

 C5 0 0 0 0 0 2.29 0 0 0.03 2.29 0.15 15.49 0 0 0 0 0 0 0 0 0 0 0 

 C6 0 0 0 0 0 2.46 0 0 0.03 2.46 0.17 16.65 0 0 0 0 0 0 0 0 0 0 0 

 C7 0 0 0 0 0 1.15 0 0 0.02 1.15 0.08 17.54 0 0 0 0 0 0 0 0 0 0 0 

 C8 0 0 0 0 0 1.18 0 0 0.02 1.18 0.08 17.99 0 0 112.62 0 0 0 0 0 0 0 0 

 C9 0 0 0 0 0 1.11 0 0 0.02 1.19 0.08 18.10 0 0 0 0 0 0 0 0 0 0 0 

 C10 0 0 0 0 0 1.16 0 0 0.02 1.19 0.08 18.12 0 0 0 0 0 0 0 0 0 0 0  C11 0 0 0 0 0 0.83 0 0 0.02 1.16 0.08 17.62 0 0 0 0 0 0 0 0 0 0 0  C12 0 0 0 0 0 0.13 0 0 0.01 0.54 0.04 17.25 0 0 0 0 0 0 0 0 0 0 0  C13 0 0 0 0 0 0.02 0 0 0.01 0.52 0.04 16.70 0 0 0 0 0 0 0 0 0 0 0  C14 0 0 0 0 0 0 0 0 0.01 0.50 0.03 16.16 0 0 0 0 0 0 0 0 0 0 0  C15 0 0 0 0 0 0 0 0 0.01 0.49 0.03 15.57 0 0 0 0 0 0 0 0 0 0 0  C16 0 0 0 0 0 0 0 0 0.01 0.47 0.03 14.93 0 0 0 266.11 0 0 0 0 0 0 0  C17 0 0 0 0 0 0 0 0 0 0 0 14.75 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 14.05 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 13.35 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 12.64 0 0 0 0 0 0 0 0 0 0 0  COS 0 0 0 0 0.30 1.20 0 0 0.02 1.12 0.08 0 0 0 0 0 0 0.13 0 0 0 0 0  AR 0 0 0 0 43.70 544.37 0 0 6.87 502.93 33.99 0 6.87 0 0 0 0 1.03 0 0 33.99 37.28 0 4 BIOMASS 2000 2000 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 00000000000000000000000 SOOT 0 0 0 0 6.00 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SLAG 0 0 0 114.00 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, soot, or slag
140 

Table E-2. Preprocessing Area Stream Data for HT Scenario 

HT A100

CH00BMASCH02BMASCH03BMASCH90BMASDR05BMASDR81STMDR84STMDR92WATGR06BMASGR08BMASGR90BMAS

Temperature (C) 25 25 25 25 90 200 120 120 90 90 90Pressure (bar) 1.01 1.01 1.01 1.01 1.01 1.98 1.98 1.98 1.01 1.01 1.01Vapor Fraction 0.00 0.00 0.00 0.00 0.00 1.00 1.00 1.00 0.00 0.00 0.00

Volume Flow** (m3/sec) 0.43 0.43 0.43 0 0.14 50.53 42.34 4.70 0.14 0.14 0

Mole Flow** (kmol/hr) 1542 1542 1542 0 513.97 9251 9251 1028 513.97 513.97 0Mass Flow (tonnes/day) 2667 2709 2667 42.20 2222 4000 4000 444.44 2240 2222 17.60 H2O 666.67 666.67 666.67 0 222.22 0 0 444.44 222.22 222.22 0 CO 0 0 0 0 0 0 0 0 0 0 0 H2 0 0 0 0 0 0 0 0 0 0 0 CO2 0 0 0 0 0 0 0 0 0 0 0 O2 0 0 0 0 0 0 0 0 0 0 0 N2 0 0 0 0 0 0 0 0 0 0 0 CH4 0 0 0 0 0 0 0 0 0 0 0 C2H6 0 0 0 0 0 0 0 0 0 0 0 C2H4 0 0 0 0 0 0 0 0 0 0 0 C2H2 0 0 0 0 0 0 0 0 0 0 0 C3 0 0 0 0 0 0 0 0 0 0 0 C4 0 0 0 0 0 0 0 0 0 0 0 H2S 0 0 0 0 0 0 0 0 0 0 0 NH3 0 0 0 0 0 0 0 0 0 0 0 TAR 0 0 0 0 0 0 0 0 0 0 0 SULFUR 0 0 0 0 0 0 0 0 0 0 0 CARBON 0 0 0 0 0 0 0 0 0 0 0 CHAR 0 0 0 0 0 0 0 0 0 0 0 STEAM 0 0 0 0 0 4000 4000 0 0 0 0 SO2 0 0 0 0 0 0 0 0 0 0 0 NO2 0 0 0 0 0 0 0 0 0 0 0 MEA 0 0 0 0 0 0 0 0 0 0 0 AIR 0 0 0 0 0 0 0 0 0 0 0 WAXES 0 0 0 0 0 0 0 0 0 0 0 C5 0 0 0 0 0 0 0 0 0 0 0 C6 0 0 0 0 0 0 0 0 0 0 0 C7 0 0 0 0 0 0 0 0 0 0 0 C8 0 0 0 0 0 0 0 0 0 0 0 C9 0 0 0 0 0 0 0 0 0 0 0 C10 0 0 0 0 0 0 0 0 0 0 0 C11 0 0 0 0 0 0 0 0 0 0 0 C12 0 0 0 0 0 0 0 0 0 0 0 C13 0 0 0 0 0 0 0 0 0 0 0 C14 0 0 0 0 0 0 0 0 0 0 0 C15 0 0 0 0 0 0 0 0 0 0 0 C16 0 0 0 0 0 0 0 0 0 0 0 C17 0 0 0 0 0 0 0 0 0 0 0 C18 0 0 0 0 0 0 0 0 0 0 0 C19 0 0 0 0 0 0 0 0 0 0 0 C20 0 0 0 0 0 0 0 0 0 0 0 COS 0 0 0 0 0 0 0 0 0 0 0 AR 0 0 0 0 0 0 0 0 0 0 0 BIOMASS 2000 2042 2000 42.20 2000 0 0 0 2018 2000 17.60 ASH 0 0 0 0 0 0 0 0 0 0 0 SOOT 0 0 0 0 0 0 0 0 0 0 0 SLAG 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, soot, or slag
141 

Table E-3. Gasification Area Stream Data for HT Scenario 

HT A200

GS02SLAGGS08BMASGS09CAIR GS11BMASGS42CO2 GS52FLUE GS54SGASGS81STMGS84STMGS90OXGS92CO2GS98STMS06SGASSL01SLAGSL02SGA

Temperature (C) 50 90 25 90 250 220 45 200 120 149 92 200 203 1300 Pressure (bar) 26.62 1.01 1.01 1.01 28.00 1.03 23.58 1.98 1.98 28.00 28.00 28.00 25.93 26.62 Vapor Fraction 0.00 0.00 1.00 0.53 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 0.00 Volume Flow** (m3/sec) 0 0.14 2.33 30.59 0.07 4.33 0.02 50.53 42.34 0.34 0.04 0.02 3.50 0 

Mole Flow** (kmol/hr) 0 513.97 343.36 6974 175.05 393.42 71.82 9251 9251 957.07 158.25 2220 8308 0 Mass Flow (tonnes/day) 114.00 2222 237.82 2222 180.00 266.91 29.08 4000 4000 743.70 166.90 960.00 3869 114.00  H2O 0 222.22 0 222.22 2.50 27.89 0 0 0 0 0 960.00 1038 0 9 CO 0 0 0 0 0.47 0 11.24 0 0 0 0.45 0 1457 0  H2 0 0 0 101.20 0 0 1.83 0 0 0 0 0 122.88 0 1 CO2 0 0 0 0 175.21 40.00 2.40 0 0 0 166.45 0 1184 0  O2 0 0 55.97 812.60 0 10.28 0 0 0 700.00 0 0 0 0  N2 0 0 181.85 16.00 1.78 181.85 0 0 0 0 0 0 17.68 0  CH4 0 0 0 0 0.01 0 0.87 0 0 0 0 0 0.02 0  C2H6 0 0 0 0 0.01 0 1.46 0 0 0 0 0 0 0  C2H4 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C2H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C3 0 0 0 0 0.01 0 1.93 0 0 0 0 0 0 0  C4 0 0 0 0 0.02 0 2.29 0 0 0 0 0 0 0  H2S 0 0 0 0 0 0 0 0 0 0 0 0 4.50 0  NH3 0 0 0 0 0 0 0 0 0 0 0 0 0.11 0  TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 4.40 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 945.60 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 4000 4000 0 0000

 SO2 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 00000000000000 C5 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0  C6 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0  C7 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0  C8 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0  C9 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0  C10 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0  C11 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0  C12 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0  C13 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0  C14 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0  C15 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0  C16 0 0 0 0 0 0 0.01 0 0 0 0 0 0 0  C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0  COS 0 0 0 0 0 0 0.02 0 0 0 0 0 0.30 0  AR 0 0 0 0 0 6.87 6.87 0 0 43.70 0 0 43.70 0  BIOMASS 0 2000 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 120.00 0 0 0 0 0 0 0 0 0 0  SOOT 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SLAG 114.00 0 0 0 0 0 0 0 0 0 0 0 0 114.00 **Volumetric and Mole flow values do not include biomass, ash, soot, or slag
142 

Table E-4. Syngas Cleaning Area Stream Data for HT Scenario 

HT A300

CL25ACG CL25SGAS CL26SGAS CL28SGAS CL29SGAS CL32SGASCL40CO2 CL41CO2 CL42CO2 CL49SGAS CL81WAT CL82WATCL83SUL CL83WAT CL90STMCL92AIRCL95AIR GS84STMSG91STMSL01H2O SL02SGA

Temperature (C) 50 203 291 240 60 40 53 53 250 45 30 40 50 60 190 25 50 62 250 203 Pressure (bar) 3.45 25.93 24.82 24.82 24.82 24.82 3.45 3.45 28.00 23.58 24.82 24.82 1.93 24.82 10.00 1.01 1.93 22.75 25.86 26.62 

Vapor Fraction 1.00 1.00 1.00 1.00 0.69 1.00 1.00 1.00 1.00 1.00 0.00 0.00 0.00 0.00 1.00 1.00 0.93 1.00 1.00 0.00 Volume Flow** (m3/sec) 3.69 1.08 2.01 4.54 2.07 1.86 3.70 0.38 0.07 1.64 0 0.02 0 0.02 1.30 0.01 0.01 3.45 0.54 0.07 

Mole Flow** (kmol/hr) 1728 2576 3847 9579 9579 6419 1716 175.05 175.05 5260 115.64 3276 13.46 2313 1271 2.07 2.58 10089 1271 9251 Mass Flow (tonnes/day) 1773 1199 1749 4418 4418 2967 1765 180.00 180.00 2130 50.00 1501 7.20 1000 549.62 1.59 2.32 3377 549.62 4000  H2O 27.18 321.91 591.14 1308 1308 9.19 24.46 2.50 2.50 0 50.00 1348 4.04 1000 549.62 0 0.45 45.53 549.62 4000 9 CO 4.62 451.69 15.74 1021 1021 999.01 4.62 0.47 0.47 823.54 0 22.10 0 0 0 0 0 1818 0 0  H2 0 38.09 69.47 154.26 154.26 154.26 0 0 0 134.08 0 0 0 0 0 0 0 288.34 0 0 1 CO2 1718 367.08 1052 1869 1869 1741 1718 175.21 175.21 175.74 0 128.26 0 0 0 0 0 190.01 0 0  O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 1.59 0.02 0 0 0  N2 17.41 5.48 5.48 17.68 17.68 17.43 17.41 1.78 1.78 0 0 0.26 0 0 0 0 0 0 0 0  CH4 0.08 0.01 0.01 0.02 0.02 0.02 0.08 0.01 0.01 63.47 0 0 0 0 0 0 0 63.41 0 0  C2H6 0.11 0 0 0 0 0 0.11 0.01 0.01 107.01 0 0 0 0 0 0 0 106.90 0 0  C2H4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C2H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C3 0.11 0 0 0 0 0 0.11 0.01 0.01 141.56 0 0 0 0 0 0 0 141.45 0 0  C4 0.17 0 0 0 0 0 0.17 0.02 0.02 167.55 0 0 0 0 0 0 0 167.38 0 0  H2S 3.39 1.40 1.40 4.50 4.50 3.48 0 0 0 0 0 1.02 0 0 0 0 0.03 0.03 0 0  NH3 0 0.03 0.03 0.11 0.11 0.01 0 0 0 0 0 0.10 0 0 0 0 0 0 0 0  TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 3.16 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 00000000000000000 C5 0 0 0 0 0 0 0 0 0 2.29 0 0 0 0 0 0 0 2.29 0 0  C6 0 0 0 0 0 0 0 0 0 2.46 0 0 0 0 0 0 0 2.46 0 0  C7 0 0 0 0 0 0 0 0 0 1.15 0 0 0 0 0 0 0 1.15 0 0  C8 0 0 0 0 0 0 0 0 0 1.18 0 0 0 0 0 0 0 1.18 0 0  C9 0.08 0 0 0 0 0 0 0 0 1.19 0 0 0 0 0 0 0.08 1.11 0 0  C10 0.03 0 0 0 0 0 0 0 0 1.19 0 0 0 0 0 0 0.03 1.16 0 0  C11 0.23 0 0 0 0 0 0 0 0 1.16 0 0 0 0 0 0 0.23 0.83 0 0  C12 0.13 0 0 0 0 0 0 0 0 0.54 0 0 0 0 0 0 0.13 0.13 0 0  C13 0.04 0 0 0 0 0 0 0 0 0.52 0 0 0 0 0 0 0.04 0.02 0 0  C14 0.01 0 0 0 0 0 0 0 0 0.50 0 0 0 0 0 0 0.01 0 0 0  C15 0 0 0 0 0 0 0 0 0 0.49 0 0 0 0 0 0 0 0 0 0  C16 0 0 0 0 0 0 0 0 0 0.47 0 0 0 0 0 0 0 0 0 0  C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  COS 0.08 0.09 0.09 0.30 0.30 0.17 0 0 0 1.12 0 0.13 0 0 0 0 0.08 1.20 0 0  AR 1.22 13.55 13.55 43.70 43.70 42.66 0 0 0 502.93 0 1.03 0 0 0 0 1.22 544.37 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SOOT 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SLAG 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, soot, or slag
143 

Table E-5. Acid Gas Removal and Sulfur Recovery Areas Stream Data for HT Scenario 

HT A300AGR

AG02MEAR AG03MEAR AG04MEAR AG06MEAR AG08MEAL AG09MEAL AG10MEAL AG11MEAL AG23ACGAG24ACGAG25ACG AG31SGAS AG32SGAS AG33SGAS AG49SGAS A300SUL SU25ACGSU26SUL SU40CO2 SU83SULSU92

Temperature (C) 58 62 58 86 50 123 96 96 86 50 50 40 58 62 45 50 53 53 50 

Pressure (bar) 22.75 22.75 22.75 3.45 3.45 3.45 3.45 20.68 3.45 3.45 3.45 24.82 22.75 22.75 23.58 3.45 3.45 3.45 1.93 

Vapor Fraction 0.00 0.00 0.00 0.01 0.00 0.00 0.00 0.00 0.98 0.83 1.00 1.00 0.93 1.00 1.00 1.00 0.49 1.00 0.00 

Volume Flow** (m3/sec) 0.480 0 0.480 1.590 0 0.480 0.470 0.470 4.820 3.690 3.690 1.860 3.410 3.450 1.640 3.690 0.010 3.700 0 

Mole Flow** (kmol/hr) 84877 699.000 85576 85576 342.110 83848 83848 83848 2070 2070 1728 6419 10788 10089 5260 1728 11.910 1716 13.460 

Mass Flow (tonnes/day) 42560 308.340 42868 42868 155.170 41096 41096 41096 1928 1928 1773 2967 3685 3377 2130 1773 7.930 1765 7.200 

 H2O 33956 297.460 34254 34254 144.090 34226 34226 34226 171.270 171.270 27.180 9.190 342.990 45.530 0 27.180 2.720 24.460 4.040 

 CO 0 4.640 4.640 4.640 0.010 0.010 0.010 0.010 4.630 4.630 4.620 999.010 1823 1818 823.540 4.620 0 4.620 0

 H2 0 0 0 0 0 0 0 0 0 0 0 154.260 288.340 288.340 134.080 0 0 0 0

 CO2 1725 1.650 1727 1727 8.690 8.690 8.690 8.690 1727 1727 1718 1741 191.660 190.010 175.740 1718 0 1718 0

 O2 000000000000000 0 0 0 0 

 N2 17.430 0 17.430 17.430 0.020 0.020 0.020 0.020 17.430 17.430 17.410 17.430 0 0 0 17.410 0 17.410 0  CH4 0 0.080 0.080 0.080 0 0 0 0 0.080 0.080 0.080 0.020 63.490 63.410 63.470 0.080 0 0.080 0  C2H6 0 0.110 0.110 0.110 0 0 0 0 0.110 0.110 0.110 0 107.010 106.900 107.010 0.110 0 0.110 0 C2H4 000000000000000 0000

 C2H2 000000000000000 0000

 C3 0 0.110 0.110 0.110 0 0 0 0 0.110 0.110 0.110 0 141.560 141.450 141.560 0.110 0 0.110 0 C4 0 0.170 0.170 0.170 0 0 0 0 0.170 0.170 0.170 0 167.550 167.380 167.550 0.170 0 0.170 0 H2S 3.450 0 3.450 3.450 0.060 0.060 0.060 0.060 3.450 3.450 3.390 3.480 0.030 0.030 0 3.390 3.390 0 0  NH3 0.010 0 0.010 0.010 0.010 0.010 0.010 0.010 0.010 0.010 0 0.010 0 0 0 0000 TAR 000000000000000 0000

 SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 3.160  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0000 CHAR 000000000000000 0000

 STEAM 000000000000000 0000

 SO2 000000000000000 0000

 NO2 000000000000000 0000

 MEA 6858 0 6858 6858 0 6858 6858 6858 0000000 0000

 AIR 000000000000000 0000

 WAXES 0 0 0 000000000000 0000

 C5 0 0 0 0 0 0 0 0 0 0 0 0 2.290 2.290 2.290 0000 C6 0 0 0 0 0 0 0 0 0 0 0 0 2.460 2.460 2.460 0000 C7 0 0 0 0 0 0 0 0 0 0 0 0 1.150 1.150 1.150 0000 C8 0 0 0 0 0 0 0 0 0 0 0 0 1.180 1.180 1.180 0000 C9 0 0.080 0.080 0.080 0.010 0.010 0.010 0.010 0.080 0.080 0.080 0 1.190 1.110 1.190 0.080 0.080 0 0  C10 0 0.030 0.030 0.030 0 0 0 0 0.030 0.030 0.030 0 1.190 1.160 1.190 0.030 0.030 0 0  C11 0 0.330 0.330 0.330 0.100 0.100 0.100 0.100 0.320 0.320 0.230 0 1.160 0.830 1.160 0.230 0.230 0 0  C12 0 0.410 0.410 0.410 0.280 0.280 0.280 0.280 0.410 0.410 0.130 0 0.540 0.130 0.540 0.130 0.130 0 0  C13 0 0.500 0.500 0.500 0.460 0.460 0.460 0.460 0.500 0.500 0.040 0 0.520 0.020 0.520 0.040 0.040 0 0  C14 0 0.500 0.500 0.500 0.490 0.490 0.490 0.490 0.500 0.500 0.010 0 0.500 0 0.500 0.010 0.010 0 0  C15 0 0.490 0.490 0.490 0.490 0.490 0.490 0.490 0.490 0.490 0 0 0.490 0 0.490 0000 C16 0 0.470 0.470 0.470 0.470 0.470 0.470 0.470 0.470 0.470 0 0 0.470 0 0.470 0000 C17 000000000000000 0000

 C18 000000000000000 0000

 C19 000000000000000 0000

 C20 000000000000000 0000

 COS 0 0.080 0.080 0.080 0 0 0 0 0.080 0.080 0.080 0.170 1.290 1.200 1.120 0.080 0.080 0 0  AR 0 1.220 1.220 1.220 0 0 0 0 1.220 1.220 1.220 42.660 545.590 544.370 502.930 1.220 1.220 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0000 ASH 000000000000000 0000

 SOOT 000000000000000 0000 SLAG 000000000000000 0000

**Volumetric and Mole flow values do not include biomass, ash, soot, or slag
144 

Table E-6. Fuel Synthesis Area Stream Data for HT Scenario 

HT A400

FS34SGAS FS35SGAS FS36SGAS FS37SGAS FS40SGAS FS42SGAS FS43SGASFS44FT FS44SGASFS45FT FS45SGASFS46FT FS46SGAS FS47SGAS FS48SGAS FS49SGASFS50FT FS50SGAS FS51SGAS FS52SGAS FS60WAT 

Temperature (C) 62 76 200 200 30 30 417 200 202 35 202 43 0 45 45 45 35 45 50 200 Pressure (bar) 22.75 26.00 26.00 24.97 1.00 1.00 24.97 23.58 24.97 23.58 24.96 23.58 0.00 23.58 23.58 23.58 22.20 23.58 24.96 24.96 22Vapor Fraction 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.86 1.00 0.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 0

Volume Flow** (m3/sec) 3.45 3.16 4.29 4.47 1.26 0.63 0.06 4.84 4.45 2.71 5.86 0.02 2.74 0.02 1.64 0.11 0.01 1.00 0.96 1.42 0

Mole Flow** (kmol/hr) 10089 10089 10089 10089 179.78 89.51 89.51 10438 9998 10438 13198 1575 8887 71.82 5260 355.46 89.49 3199 3199 3199 14Mass Flow (tonnes/day) 3377 3377 3377 3377 60.18 55.81 55.81 4668 3372 4668 4668 1069 3599 29.08 2130 143.95 427.14 1296 1296 1296 641 H2O 45.53 45.53 45.53 45.53 0.81 0.81 0.81 642.07 45.53 642.07 45.53 642.07 0 0 0 0 0 0 0 0 641

 CO 1818 1818 1818 1818 32.40 32.40 32.40 1391 1818 1391 2319 0 1391 11.24 823.54 55.65 0 500.87 500.87 500.87  H2 288.34 288.34 288.34 288.34 5.14 0.77 0.77 225.34 283.97 225.34 365.52 0 226.52 1.83 134.08 9.06 0 81.55 81.55 81.55 

 CO2 190.01 190.01 190.01 190.01 3.39 3.39 3.39 296.89 190.01 296.89 296.89 0 296.90 2.40 175.74 11.88 0 106.88 106.88 106.88  O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CH4 63.41 63.41 63.41 63.41 1.13 1.13 1.13 107.33 63.41 107.33 102.02 0 107.23 0.87 63.47 4.29 0 38.60 38.60 38.60  C2H6 106.90 106.90 106.90 106.90 1.91 1.91 1.91 180.95 106.90 180.95 171.99 0 180.79 1.46 107.01 7.23 0 65.08 65.08 65.08  C2H4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C2H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C3 141.45 141.45 141.45 141.45 2.52 2.52 2.52 239.37 141.45 239.37 227.54 0 239.16 1.93 141.56 9.57 0 86.10 86.10 86.10  C4 167.38 167.38 167.38 167.38 2.98 2.98 2.98 283.31 167.38 283.31 269.28 0 283.06 2.29 167.55 11.32 0 101.90 101.90 101.90  H2S 0.03 0.03 0.03 0 0 0 0 0 0 0 0 0 0.00 0 0 0 0 0 0 0  NH3 0 0 0 0 0 0 0 0 0 0 0 0 0.00 0 0 0 0 0 0 0  TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 170.22 0 170.22 0 170.22 0 0 0 0 170.22 0 0 0 

 C5 2.29 2.29 2.29 2.29 0.04 0.04 0.04 19.36 2.29 19.36 3.69 15.49 3.87 0.03 2.29 0.15 15.49 1.39 1.39 1.39 

 C6 2.46 2.46 2.46 2.46 0.04 0.04 0.04 20.81 2.46 20.81 3.96 16.65 4.16 0.03 2.46 0.17 16.65 1.50 1.50 1.50 

 C7 1.15 1.15 1.15 1.15 0.02 0.02 0.02 19.49 1.15 19.49 1.85 17.54 1.95 0.02 1.15 0.08 17.54 0.70 0.70 0.70 

 C8 1.18 1.18 1.18 1.18 0.02 0.02 0.02 19.99 1.18 19.99 1.90 17.99 2.00 0.02 1.18 0.08 17.99 0.72 0.72 0.72 

 C9 1.11 1.11 1.11 1.11 0.02 0.02 0.02 20.11 1.11 20.11 1.83 18.10 2.01 0.02 1.19 0.08 18.10 0.72 0.72 0.72 

 C10 1.16 1.16 1.16 1.16 0.02 0.02 0.02 20.13 1.16 20.13 1.88 18.12 2.01 0.02 1.19 0.08 18.12 0.72 0.72 0.72  C11 0.83 0.83 0.83 0.83 0.01 0.01 0.01 19.58 0.83 19.58 1.54 17.62 1.96 0.02 1.16 0.08 17.62 0.71 0.71 0.71  C12 0.13 0.13 0.13 0.13 0 0 0 18.16 0.13 18.16 0.45 17.25 0.91 0.01 0.54 0.04 17.25 0.33 0.33 0.33  C13 0.02 0.02 0.02 0.02 0 0 0 17.58 0.02 17.58 0.33 16.70 0.88 0.01 0.52 0.04 16.70 0.32 0.32 0.32  C14 0 0 0 0 0 0 0 17.01 0 17.01 0.31 16.16 0.85 0.01 0.50 0.03 16.16 0.31 0.31 0.31  C15 0 0 0 0 0 0 0 16.39 0 16.39 0.30 15.57 0.82 0.01 0.49 0.03 15.57 0.29 0.29 0.29  C16 0 0 0 0 0 0 0 15.72 0 15.72 0.28 14.93 0.79 0.01 0.47 0.03 14.93 0.28 0.28 0.28  C17 0 0 0 0 0 0 0 14.75 0 14.75 0 14.75 0 0 0 0 14.75 0 0 0  C18 0 0 0 0 0 0 0 14.05 0 14.05 0 14.05 0 0 0 0 14.05 0 0 0  C19 0 0 0 0 0 0 0 13.35 0 13.35 0 13.35 0 0 0 0 13.35 0 0 0  C20 0 0 0 0 0 0 0 12.64 0 12.64 0 12.64 0 0 0 0 12.64 0 0 0  COS 1.20 1.20 1.20 1.20 0.02 0.02 0.02 1.88 1.20 1.88 1.88 0 1.88 0.02 1.12 0.08 0 0.68 0.68 0.68  AR 544.37 544.37 544.37 544.37 9.70 9.70 9.70 850.25 544.37 850.25 850.25 0 849.66 6.87 502.93 33.99 0 305.88 305.88 305.88  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SOOT 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SLAG 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, soot, or slag
145 

Table E-7. Hydroprocessing, Power Generation, and Air Separation Areas Stream Data for HT Sce

HT A500 HY50FT HY65FGAS HY70GASO HY71GASO HY80DIES HY81DIES HY90HYD HT A600 421 602 615 680 CW2CWRECHPSTM1HPSTM2 HT A700 AIR-1A AIR-A AIR-LN2-OUTemperature (C) 35 35 35 36 35 37 30 45 1144 273 30 85 73 170 565 -170 32 20 Pressure (bar) 22.20 22.20 22.20 1.03 22.20 1.03 1.00 23.58 1.00 1.00 1.01 7.91 0.30 7.91 173.38 6.20 1.01 6.30 

Vapor Fraction 0.00 1.00 0.00 0.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00 0.00 0.97 0.00 0.00 1.00 1.00 1.00 

Volume Flow** (m3/sec) 0.01 0.02 0 0 0.01 0.01 0.63 0.11 117.68 45.37 22.37 0.01 49.13 0.01 0.01 1.33 29.07 4.49 

Mole Flow** (kmol/hr) 89.49 74.55 41.08 41.08 48.97 48.97 90.27 355.46 3595 3595 3237 1939 1939 1939 1939 3968 4177 4177 Mass Flow (tonnes/day) 427.14 52.77 112.62 112.62 266.11 266.11 4.37 143.95 2439 2439 2242 838.24 838.23 838.23 838.23 2758 2903 2903 

 H2O 0 0 0 0 0 0 0 0 233.43 233.43 0 838.24 838.23 838.23 838.23 0 0 0  CO 0 0 0 0 0 0 0 55.65 0 0 0 0 0 0 0 0 0 0  H2 0 0 0 0 0 0 4.37 9.06 0 0 0 0 0 0 0 0 0 0  CO2 0 0 0 0 0 0 0 11.88 352.24 352.24 0 0 0 0 0 0 0 0  O2 0 0 0 0 0 0 0 0 104.25 104.25 527.27 0 0 0 0 638.58 672.19 672.19  N2 0 0 0 0 0 0 0 0 1715 1715 1715 0 0 0 0 2084 2194 2194 

 CH4 0 14.94 0 0 0 0 0 4.29 0 0 0 0 0 0 0 0 0 0  C2H6 0 0 0 0 0 0 0 7.23 0 0 0 0 0 0 0 0 0 0  C2H4 0000000 00000000 000 C2H2 0000000 00000000 000 C3 0 37.83 0 0 0 0 0 9.57 0 0 0 0 0 0 0 0 0 0  C4 0 0 0 0 0 0 0 11.32 0 0 0 0 0 0 0 0 0 0  H2S 0000000 00000000 000 NH3 0000000 00000000 000 TAR 0000000 00000000 000 SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0000000 00000000 000 STEAM 0000000 00000000 000 SO2 0 0 0 0 0 0 0 0 0.08 0.08 0 0 0 0 0 0 0 0  NO2 0000000 00000000 000 MEA 0000000 00000000 000 AIR 0000000 00000000 000 WAXES 170.22000000 00000000 000 C5 15.49 0 0 0 0 0 0 0.15 0 0 0 0 0 0 0 0 0 0  C6 16.65 0 0 0 0 0 0 0.17 0 0 0 0 0 0 0 0 0 0  C7 17.54 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0  C8 17.99 0 112.62 112.62 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0  C9 18.10 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0  C10 18.12 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0  C11 17.62 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0  C12 17.25 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C13 16.70 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C14 16.16 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0 0 0  C15 15.57 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0 0 0  C16 14.93 0 0 0 266.11 266.11 0 0.03 0 0 0 0 0 0 0 0 0 0  C17 14.75 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 14.05 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 13.35 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 12.64 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  COS 0 0 0 0 0 0 0 0.08 0 0 0 0 0 0 0 0 0 0  AR 0 0 0 0 0 0 0 33.99 33.99 33.99 0 0 0 0 0 35.42 37.28 37.28  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0000000 00000000 000 SOOT 0000000 00000000 000 SLAG 0000000 00000000 000**Volumetric and Mole flow values do not include biomass, ash, soot, or slag
146 

Low-Temperature Scenario 

Table E-8. Overall Plant Stream Data for LT Scenario 

HT Overall Plant

PL00BMAS PL08BMAS PL09CAIR PL21SGAS PL34SGASPL41CO2 PL43CO2 PL52FLUE PL55SGASPL56FT PL56SGASPL60ASH PL65FGAS PL71GASOPL81DIESPL81STMPL81WATPL83SUL PL84STMPL88FLUEPL90HYDPL90OXPL90STMPL92W

Temperature (C) 25 90 100 870 62 50 243 200 32 35 32 0 50 50 50 200 40 50 120 344 60 149 300 Pressure (bar) 1.01 1.01 1.01 27.55 22.89 3.45 28.00 1.00 22.89 22.89 22.89 1.00 22.89 22.89 22.89 1.98 26.86 2.07 1.98 1.00 1.01 22.00 25.00 

Vapor Fraction 0.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 1.00 0.00 1.00 0.00 0.00 1.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00 Volume Flow** (m3/sec) 0.43 0.14 16.89 4.76 2.40 2.84 0.07 22.48 1.50 0.01 0.12 0 0.02 0 0 50.53 0.02 0 42.34 40.84 0.53 0.32 1.15 

Mole Flow** (kmol/hr) 1542 513.97 1985 4939 7066 1334 170.42 2053 4869 69.40 394.75 0 57.67 31.78 37.88 9251 2804 5.56 9251 2863 69.96 722.81 2313 

Mass Flow (tonnes/day) 2667 2222 1375 2930 2706 1380 180.00 1471 2071 330.42 167.90 118.88 40.83 87.12 205.86 4000 1388 3.10 4000 1955 3.38 561.66 1000  H2O 666.67 222.22 0 413.42 30.18 20.50 0 27.16 0 0.09 0 0 0 0 0 0 1060 2.40 0 189.23 0 0 1000  CO 0 0 0 797.86 1575 0 0 0 795.10 0 64.47 0 0 0 0 0 15.27 0 0 0 0 0 0  H2 0 0 0 47.75 168.38 0 0 0 120.71 0 9.79 0 0 0 0 0 0.07 0 0 0 3.38 0 0  CO2 0 0 0 1427 167.80 1359 180.00 239.52 559.11 0 45.33 0 0 0 0 0 263.09 0 0 318.35 0 0 0  O2 0 0 320.19 0 0 0 0 148.99 0 0 0 0 0 0 0 0 0 0 0 80.99 0 528.66 0  N2 0 0 1055 0 0 0 0 1055 0 0 0 0 0 0 0 0 0 0 0 1336 0 0 0  AR 0 0 0 33.00 411.50 0 0 0 380.61 0 30.86 0 0 0 0 0 0.69 0 0 30.86 0 33.00 0  CH4 0 0 0 103.81 149.06 0 0 0 52.01 0 4.22 0 11.56 0 0 0 4.58 0 0 0 0 0 0  C2H6 0 0 0 21.82 28.49 0 0 0 13.55 0 1.10 0 0 0 0 0 5.75 0 0 0 0 0 0  C2H4 0 0 0 44.96 36.05 0 0 0 0.36 0 0.03 0 0 0 0 0 8.11 0 0 0 0 0 0  C6H6 0 0 0 4.10 39.50 0 0 0 36.54 0 2.96 0 0 0 0 0 0.98 0 0 0 0 0 0  C3 0 0 0 0 54.18 0 0 0 58.62 0 4.75 0 29.27 0 0 0 0 0 0 0 0 0 0  C4 0 0 0 0 34.31 0 0 0 41.73 0 3.38 0 0 0 0 0 0 0 0 0 0 0 0  H2S 0 0 0 4.51 0.13 0 0 0 0 0 0 0 0 0 0 0 1.87 0 0 0 0 0 0  NH3 0 0 0 19.09 1.52 0 0 0 0.28 0 0.02 0 0 0 0 0 17.36 0 0 0 0 0 0  TAR 0 0 0 10.57 0 0 0 0 0 0 0 0 0 0 0 0 10.57 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 4000 0 0 4000 0 0 0 0  SO2 0 0 0 0 0 0 0 0.30 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.06 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 0 0 131.46 0 0 0 0 0 0 0 0 0 0 0 0 0  C5 0 0 0 0 2.21 0 0 0 2.21 11.92 0.18 0 0 0 0 0 0 0 0 0 0 0 0  C6 0 0 0 0 2.37 0 0 0 2.37 12.82 0.19 0 0 0 0 0 0 0 0 0 0 0 0  C7 0 0 0 0 1.11 0 0 0 1.11 13.53 0.09 0 0 0 0 0 0 0 0 0 0 0 0  C8 0 0 0 0 1.14 0 0 0 1.14 13.88 0.09 0 0 87.12 0 0 0 0 0 0 0 0 0  C9 0 0 0 0 1.08 0 0 0 1.15 13.95 0.09 0 0 0 0 0 0 0 0 0 0 0 0  C10 0 0 0 0 1.13 0 0 0 1.15 13.98 0.09 0 0 0 0 0 0 0 0 0 0 0 0  C11 0 0 0 0 0.50 0 0 0 0.54 13.86 0.04 0 0 0 0 0 0 0 0 0 0 0 0  C12 0 0 0 0 0.44 0 0 0 0.53 13.54 0.04 0 0 0 0 0 0 0 0 0 0 0 0  C13 0 0 0 0 0.34 0 0 0 0.51 13.10 0.04 0 0 0 0 0 0 0 0 0 0 0 0  C14 0 0 0 0 0.20 0 0 0 0.49 12.57 0.04 0 0 0 0 0 0 0 0 0 0 0 0  C15 0 0 0 0 0 0 0 0 0.46 11.93 0.04 0 0 0 0 0 0 0 0 0 0 0 0  C16 0 0 0 0 0.04 0 0 0 0.45 11.48 0.04 0 0 0 205.86 0 0 0 0 0 0 0 0  C17 0 0 0 0 0 0 0 0 0 11.39 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 10.85 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 10.31 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 9.76 0 0 0 0 0 0 0 0 0 0 0 0 0  BIOMASS 2000 2000 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 1.01 0 0 0 0.01 0 0 0 118.88 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0.81 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
147 

Table E-9. Preprocessing Area Stream Data for LT Scenario 

HT A100 CH00BMAS CH02BMAS CH03BMAS CH90BMAS DR05BMAS .DR81STM.DR84STM.DR92WAT GR06BMAS GR08BMAS GR90BMASTemperature (C) 25 25 25 0 90 200 120 120 90 90 0Pressure (bar) 1.01 1.01 1.01 0.00 1.01 1.98 1.98 1.98 1.01 1.01 0.00Vapor Fraction 0.00 0.00 0.00 0.00 0.00 1.00 1.00 1.00 0.00 0.00 0.00Volume Flow** (m3/sec) 0.43 0.43 0.43 0 0.14 50.53 42.34 4.70 0.14 0.14 0

Mole Flow** (kmol/hr) 1542 1542 1542 0 513.97 9251 9251 1028 513.97 513.97 0Mass Flow (tonnes/day) 2667 3181 2667 514.02 2222 4000 4000 444.44 3019 2222 796.81 H2O 666.67 666.67 666.67 0 222.22 0 0 444.44 222.22 222.22 0 CO 0 0 0 0 0 0 0 0 0 0 0 H2 0 0 0 0 0 0 0 0 0 0 0 CO2 0 0 0 0 0 0 0 0 0 0 0 O2 0 0 0 0 0 0 0 0 0 0 0 N2 0 0 0 0 0 0 0 0 0 0 0 AR 0 0 0 0 0 0 0 0 0 0 0 CH4 0 0 0 0 0 0 0 0 0 0 0 C2H6 0 0 0 0 0 0 0 0 0 0 0 C2H4 0 0 0 0 0 0 0 0 0 0 0 C6H6 0 0 0 0 0 0 0 0 0 0 0 C3 0 0 0 0 0 0 0 0 0 0 0 C4 0 0 0 0 0 0 0 0 0 0 0 H2S 0 0 0 0 0 0 0 0 0 0 0 NH3 0 0 0 0 0 0 0 0 0 0 0 TAR 0 0 0 0 0 0 0 0 0 0 0 SULFUR 0 0 0 0 0 0 0 0 0 0 0 CARBON 0 0 0 0 0 0 0 0 0 0 0 STEAM 0 0 0 0 0 4000 4000 0 0 0 0 SO2 0 0 0 0 0 0 0 0 0 0 0 NO2 0 0 0 0 0 0 0 0 0 0 0 MEA 0 0 0 0 0 0 0 0 0 0 0 AIR 0 0 0 0 0 0 0 0 0 0 0 WAXES 0 0 0 0 0 0 0 0 0 0 0 C5 0 0 0 0 0 0 0 0 0 0 0 C6 0 0 0 0 0 0 0 0 0 0 0 C7 0 0 0 0 0 0 0 0 0 0 0 C8 0 0 0 0 0 0 0 0 0 0 0 C9 0 0 0 0 0 0 0 0 0 0 0 C10 0 0 0 0 0 0 0 0 0 0 0 C11 0 0 0 0 0 0 0 0 0 0 0 C12 0 0 0 0 0 0 0 0 0 0 0 C13 0 0 0 0 0 0 0 0 0 0 0 C14 0 0 0 0 0 0 0 0 0 0 0 C15 0 0 0 0 0 0 0 0 0 0 0 C16 0 0 0 0 0 0 0 0 0 0 0 C17 0 0 0 0 0 0 0 0 0 0 0 C18 0 0 0 0 0 0 0 0 0 0 0 C19 0 0 0 0 0 0 0 0 0 0 0 C20 0 0 0 0 0 0 0 0 0 0 0 BIOMASS 2000 2514 2000 514.02 2000 0 0 0 2797 2000 796.81 ASH 0 0 0 0 0 0 0 0 0 0 0 CHAR 0 0 0 0 0 0 0 0 0 0 0**Volumetric and Mole flow values do not include biomass, ash, or char
148 

Table E-10. Gasification Area Stream Data for LT Scenario 

HT A200

CB52FLUECB61ASHCB62ASH CY14SGAS CY15CHAR CY16CHAR GS05SGAS GS07SGAS GS09CAIR GS11BMAS GS19CHAR GS20SGAS GS43CO2 GS46CO2 GS52FLUE GS81STMGS84STMGS90O

Temperature (C) 1200 0 0 870 0 0 871 870 100 96 0 870 243 100 200 200 120 Pressure (bar) 1.00 1.00 1.00 27.57 27.57 27.55 28.00 27.58 1.01 1.01 27.55 27.55 28.00 27.58 1.00 1.98 1.98 Vapor Fraction 1.00 0.00 0.00 1.00 0.00 0.00 1.00 1.00 1.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 Volume Flow** (m3/sec) 69.99 0 0 2.38 0 0 4.69 4.76 16.89 0 0 4.76 0.07 0.05 22.48 50.53 42.34 

Mole Flow** (kmol/hr) 2053 0 0 2470 0 0 4931 4939 1985 513.97 0 4939 170.42 161.90 2053 9251 9251 7Mass Flow (tonnes/day) 1471 17.82 118.88 1477 95.38 12.09 3136 3145 1375 2222 214.95 2930 180.00 171.00 1471 4000 4000 5 H2O 27.16 0 0 206.71 0 0 413.42 413.42 0 222.22 0 413.42 0 0 27.16 0 0  CO 0 0 0 398.93 0 0 797.86 797.86 0 0 0 797.86 0 0 0 0 0  H2 0 0 0 23.87 0 0 47.75 47.75 0 0 0 47.75 0 0 0 0 0  CO2 239.52 0 0 713.66 0 0 1418 1427 0 0 0 1427 180.00 171.00 239.52 0 0  O2 148.99 0 0 0 0 0 0 0 320.19 0 0 0 0 0 148.99 0 0 5 N2 1055 0 0 0 0 0 0 0 1055 0 0 0 0 0 1055 0 0  AR 0 0 0 16.50 0 0 33.00 33.00 0 0 0 33.00 0 0 0 0 0  CH4 0 0 0 51.90 0 0 103.81 103.81 0 0 0 103.81 0 0 0 0 0  C2H6 0 0 0 10.91 0 0 21.82 21.82 0 0 0 21.82 0 0 0 0 0  C2H4 0 0 0 22.48 0 0 44.96 44.96 0 0 0 44.96 0 0 0 0 0  C6H6 0 0 0 2.05 0 0 4.10 4.10 0 0 0 4.10 0 0 0 0 0  C3 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  H2S 0 0 0 2.25 0 0 4.51 4.51 0 0 0 4.51 0 0 0 0 0  NH3 0 0 0 9.54 0 0 19.09 19.09 0 0 0 19.09 0 0 0 0 0  TAR 0 0 0 5.29 0 0 10.57 10.57 0 0 0 10.57 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 4000 4000  SO2 0.30 0 0 0 0 0 0 0 0 0 0 0 0 0 0.30 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C5 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C6 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C7 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C8 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C9 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C10 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C11 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C12 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C13 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C14 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C15 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C16 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 2000 0 0 0 0 0 0 0  ASH 0.01 17.82 118.88 7.19 52.75 6.69 119.90 119.90 0 0 118.89 1.01 0 0 0.01 0 0  CHAR 0 0 0 5.81 42.63 5.40 96.87 96.87 0 0 96.06 0.81 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
149 

Table E-11. Syngas Cleaning Area Stream Data for LT Scenario 

HT A300 CL01H2OCL02H2O CL21SGASCL25ACG CL25SGASCL25WAT CL29SGASCL40CO2 CL40SGASCL41CO2CL43CO2 CL49SGASCL80H2OCL80WATCL81WATCL83SULCL83WATCL84WTemperature (C) 194 150 870 50 194 0 40 50 62 50 243 32 30 30 40 50 98 Pressure (bar) 27.55 27.55 27.55 3.45 27.55 27.55 26.86 3.45 22.89 3.45 28.00 22.89 27.57 27.57 26.86 2.07 27.55 Vapor Fraction 0.00 0.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 0.00 0.00 0.00 0.00 Volume Flow** (m3/sec) 0.08 0.08 4.76 3.28 2.03 0 0.98 3.21 2.40 2.84 0.07 1.50 0 0.01 0.02 0 0.07 

Mole Flow** (kmol/hr) 13877 13877 4939 1538 5286 0 3638 1505 7066 1334 170.42 4869 346.93 1156 2804 5.56 13877 Mass Flow (tonnes/day) 6000 6000 2930 1586 3078 1.82 2190 1560 2706 1380 180.00 2071 150.00 500.00 1388 3.10 6000  H2O 6000 6000 413.42 22.78 563.42 0 3.78 20.50 30.18 20.50 0 0 150.00 500.00 1060 2.40 6000  CO 0 0 797.86 3.07 797.86 0 782.59 0 1575 0 0 795.10 0 0 15.27 0 0  H2 0 0 47.75 0 47.75 0 47.68 0 168.38 0 0 120.71 0 0 0.07 0 0  CO2 0 0 1427 1539 1427 0 1164 1539 167.80 1359 180.00 559.11 0 0 263.09 0 0  O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AR 0 0 33.00 1.42 33.00 0 32.31 0 411.50 0 0 380.61 0 0 0.69 0 0  CH4 0 0 103.81 2.18 103.81 0 99.23 0 149.06 0 0 52.01 0 0 4.58 0 0  C2H6 0 0 21.82 1.11 21.82 0 16.07 0 28.49 0 0 13.55 0 0 5.75 0 0  C2H4 0 0 44.96 1.14 44.96 0 36.85 0 36.05 0 0 0.36 0 0 8.11 0 0  C6H6 0 0 4.10 0.15 4.10 0 3.11 0 39.50 0 0 36.54 0 0 0.98 0 0  C3 0 0 0 4.29 0 0 0 0 54.18 0 0 58.62 0 0 0 0 0  C4 0 0 0 6.73 0 0 0 0 34.31 0 0 41.73 0 0 0 0 0  H2S 0 0 4.51 2.48 4.51 0 2.64 0 0.13 0 0 0 0 0 1.87 0 0  NH3 0 0 19.09 0.42 19.09 0 1.73 0 1.52 0 0 0.28 0 0 17.36 0 0  TAR 0 0 10.57 0 10.57 0 0 0 0 0 0 0 0 0 10.57 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C5 0 0 0 0 0 0 0 0 2.21 0 0 2.21 0 0 0 0 0  C6 0 0 0 0 0 0 0 0 2.37 0 0 2.37 0 0 0 0 0  C7 0 0 0 0 0 0 0 0 1.11 0 0 1.11 0 0 0 0 0  C8 0 0 0 0 0 0 0 0 1.14 0 0 1.14 0 0 0 0 0  C9 0 0 0 0.07 0 0 0 0 1.08 0 0 1.15 0 0 0 0 0  C10 0 0 0 0.02 0 0 0 0 1.13 0 0 1.15 0 0 0 0 0  C11 0 0 0 0.03 0 0 0 0 0.50 0 0 0.54 0 0 0 0 0  C12 0 0 0 0.07 0 0 0 0 0.44 0 0 0.53 0 0 0 0 0  C13 0 0 0 0.10 0 0 0 0 0.34 0 0 0.51 0 0 0 0 0  C14 0 0 0 0.09 0 0 0 0 0.20 0 0 0.49 0 0 0 0 0  C15 0 0 0 0 0 0 0 0 0 0 0 0.46 0 0 0 0 0  C16 0 0 0 0.02 0 0 0 0 0.04 0 0 0.45 0 0 0 0 0  C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 1.01 0 0 1.01 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0.81 0 0 0.81 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
150 

Table E-12. Acid Gas Removal and Sulfur Recovery Areas Stream Data for LT Scenario 

HT A300AGR AG02MEAR AG03MEAR AG04MEAR AG06MEAR AG08MEAL AG09MEAL AG10MEAL AG11MEAL AG23ACGAG24ACGAG25ACG AG26SGAS AG32SGAS AG33SGAS AG49SGAS A300SUL SU25ACGSU26SUL SU40CTemperature (C) 50 62 50 93 50 123 80 80 93 50 50 40 50 62 32 50 50 Pressure (bar) 22.89 22.89 22.89 3.45 3.45 3.45 3.45 26.00 3.45 3.45 3.45 26.86 22.89 22.89 22.89 3.45 3.45 

Vapor Fraction 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.00 0.88 1.00 1.00 0.93 1.00 1.00 1.00 0.82 

Volume Flow** (m3/sec) 0.31 0 0.31 0.33 0 0.31 0.30 0.30 4.24 3.28 3.28 0.98 2.30 2.40 1.50 3.28 0.06 

Mole Flow** (kmol/hr) 54067 449.42 54516 54516 207.83 52978 52978 52978 1746 1746 1538 3638 7515 7066 4869 1538 33.30 Mass Flow (tonnes/day) 28640 208.71 28848 28848 101.05 27263 27263 27263 1687 1687 1586 2190 2915 2706 2071 1586 25.65  H2O 20923 181.16 21104 21104 82.74 21081 21081 21081 105.52 105.52 22.78 3.78 211.34 30.18 0 22.78 2.28  CO 0 3.07 3.07 3.07 0 0 0 0 3.07 3.07 3.07 782.59 1578 1575 795.10 3.07 3.07  H2 0 0 0 0 0 0 0 0 0 0 0 47.68 168.38 168.38 120.71 0 0  CO2 1551 4.53 1556 1556 16.15 16.15 16.15 16.15 1556 1556 1539 1164 172.33 167.80 559.11 1539 0  O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AR 0 1.42 1.42 1.42 0 0 0 0 1.42 1.42 1.42 32.31 412.92 411.50 380.61 1.42 1.42  CH4 0 2.19 2.19 2.19 0.01 0.01 0.01 0.01 2.19 2.19 2.18 99.23 151.24 149.06 52.01 2.18 2.18  C2H6 0 1.13 1.13 1.13 0.02 0.02 0.02 0.02 1.13 1.13 1.11 16.07 29.62 28.49 13.55 1.11 1.11  C2H4 0 1.16 1.16 1.16 0.01 0.01 0.01 0.01 1.16 1.16 1.14 36.85 37.21 36.05 0.36 1.14 1.14  C6H6 0 0.15 0.15 0.15 0 0 0 0 0.15 0.15 0.15 3.11 39.65 39.50 36.54 0.15 0.15  C3 0 4.44 4.44 4.44 0.15 0.15 0.15 0.15 4.44 4.44 4.29 0 58.62 54.18 58.62 4.29 4.29  C4 0 7.42 7.42 7.42 0.69 0.69 0.69 0.69 7.41 7.41 6.73 0 41.73 34.31 41.73 6.73 6.73  H2S 2.51 0.01 2.52 2.52 0.04 0.04 0.04 0.04 2.52 2.52 2.48 2.64 0.13 0.13 0 2.48 2.48  NH3 0 0.49 0.49 0.49 0.07 0.07 0.07 0.07 0.49 0.49 0.42 1.73 2.01 1.52 0.28 0.42 0.42  TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 6163 0 6163 6163 0 6163 6163 6163 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C5 0 0 0 0 0 0 0 0 0 0 0 0 2.21 2.21 2.21 0 0  C6 0 0 0 0 0 0 0 0 0 0 0 0 2.37 2.37 2.37 0 0  C7 0 0 0 0 0 0 0 0 0 0 0 0 1.11 1.11 1.11 0 0  C8 0 0 0 0 0 0 0 0 0 0 0 0 1.14 1.14 1.14 0 0  C9 0 0.07 0.07 0.07 0 0 0 0 0.07 0.07 0.07 0 1.15 1.08 1.15 0.07 0.07  C10 0 0.02 0.02 0.02 0 0 0 0 0.02 0.02 0.02 0 1.15 1.13 1.15 0.02 0.02  C11 0 0.04 0.04 0.04 0 0 0 0 0.04 0.04 0.03 0 0.54 0.50 0.54 0.03 0.03  C12 0 0.08 0.08 0.08 0.02 0.02 0.02 0.02 0.08 0.08 0.07 0 0.53 0.44 0.53 0.07 0.07  C13 0 0.17 0.17 0.17 0.08 0.08 0.08 0.08 0.17 0.17 0.10 0 0.51 0.34 0.51 0.10 0.10  C14 0 0.29 0.29 0.29 0.20 0.20 0.20 0.20 0.29 0.29 0.09 0 0.49 0.20 0.49 0.09 0.09  C15 0 0.46 0.46 0.46 0.46 0.46 0.46 0.46 0.46 0.46 0 0 0.46 0 0.46 0 0  C16 0 0.41 0.41 0.41 0.38 0.38 0.38 0.38 0.41 0.41 0.02 0 0.45 0.04 0.45 0.02 0.02  C17 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
151 

Table E-13. Fuel Synthesis Area Stream Data for LT Scenario 

HT A400

FS34SGASFS35SGASFS36SGASFS37SGASFS40SGASFS41GASFS50SGAS FS51FT FS52FTFS53SGAS FS54FTFS55SGAS FS56FTFS56SGASFS57SGASFS58SGAS

Temperature (C) 62 78 200 200 60 235 200 200 35 32 35 32 35 32 32 4Pressure (bar) 22.89 26.50 25.12 25.12 1.01 25.00 24.96 22.89 22.89 22.89 22.89 22.89 22.89 22.89 22.89 25.3Vapor Fraction 1.00 1.00 1.00 1.00 0.88 1.00 1.00 1.00 0.72 1.00 0.00 1.00 0.00 1.00 1.00 1.0Volume Flow** (m3/sec) 2.40 2.18 4.44 0.07 0.65 0.05 5.02 4.39 2.08 2.02 0.01 1.50 0.01 0.12 0.41 0.3

Mole Flow** (kmol/hr) 7066 7066 10153 166.97 97.01 97.01 11399 9268 9268 6579 70.21 4869 69.40 394.75 1316 131Mass Flow (tonnes/day) 2706 2706 3705 60.93 57.55 57.55 4261 4261 4261 2798 330.77 2071 330.42 167.90 559.66 559.6 H2O 30.18 30.18 671.91 11.05 11.05 11.05 671.91 1133 1133 0 0.43 0 0.09 0 0  CO 1575 1575 1576 25.92 25.92 25.92 1791 1074 1074 1074 0 795.10 0 64.47 214.89 214.8 H2 168.38 168.38 242.13 3.98 0.60 0.60 271.37 163.12 163.12 163.12 0 120.71 0 9.79 32.62 32.6 CO2 167.80 167.80 604.44 9.94 9.94 9.94 755.55 755.55 755.55 755.55 0 559.11 0 45.33 151.11 151.1 O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  N2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AR 411.50 411.50 411.50 6.77 6.77 6.77 514.36 514.36 514.36 514.34 0 380.61 0 30.86 102.87 102.8 CH4 149.06 149.06 52.13 0.86 0.86 0.86 66.19 70.29 70.29 70.29 0 52.01 0 4.22 14.06 14.0 C2H6 28.49 28.49 7.72 0.13 0.13 0.13 11.38 18.30 18.30 18.30 0 13.55 0 1.10 3.66 3.6 C2H4 36.05 36.05 0.39 0.01 0.01 0.01 0.49 0.49 0.49 0.49 0 0.36 0 0.03 0.10 0.1 C6H6 39.50 39.50 39.50 0.65 0.65 0.65 49.38 49.38 49.38 49.38 0 36.54 0 2.96 9.88 9.8 C3 54.18 54.18 54.18 0.89 0.89 0.89 70.02 79.16 79.16 79.22 0 58.62 0 4.75 15.84 15.8 C4 34.31 34.31 34.31 0.56 0.56 0.56 45.59 56.43 56.43 56.40 0 41.73 0 3.38 11.28 11.2 H2S 0.13 0.13 0.01 0 0 0 0.01 0.01 0.01 0.01 0 0 0 0 0  NH3 1.52 1.52 0.30 0 0 0 0.38 0.38 0.38 0.38 0 0.28 0 0.02 0.08 0.0 TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 131.46 131.46 0 131.46 0 131.46 0 0  C5 2.21 2.21 2.21 0.04 0.04 0.04 2.80 14.91 14.91 2.98 11.92 2.21 11.92 0.18 0.60 0.6 C6 2.37 2.37 2.37 0.04 0.04 0.04 3.01 16.02 16.02 3.20 12.82 2.37 12.82 0.19 0.64 0.6 C7 1.11 1.11 1.11 0.02 0.02 0.02 1.41 15.03 15.03 1.50 13.53 1.11 13.53 0.09 0.30 0.3 C8 1.14 1.14 1.14 0.02 0.02 0.02 1.45 15.42 15.42 1.54 13.88 1.14 13.88 0.09 0.31 0.3 C9 1.08 1.08 1.08 0.02 0.02 0.02 1.39 15.50 15.50 1.55 13.95 1.15 13.95 0.09 0.31 0.3 C10 1.13 1.13 1.13 0.02 0.02 0.02 1.44 15.54 15.54 1.55 13.98 1.15 13.98 0.09 0.31 0.3 C11 0.50 0.50 0.50 0.01 0.01 0.01 0.65 14.59 14.59 0.73 13.86 0.54 13.86 0.04 0.15 0.1 C12 0.44 0.44 0.44 0.01 0.01 0.01 0.59 14.26 14.26 0.71 13.54 0.53 13.54 0.04 0.14 0.1 C13 0.34 0.34 0.34 0.01 0.01 0.01 0.47 13.79 13.79 0.69 13.10 0.51 13.10 0.04 0.14 0.1 C14 0.20 0.20 0.20 0 0 0 0.34 13.23 13.23 0.66 12.57 0.49 12.57 0.04 0.13 0.1 C15 0 0 0 0 0 0 0.13 12.55 12.55 0.63 11.93 0.46 11.93 0.04 0.13 0.1 C16 0.04 0.04 0.04 0 0 0 0.16 12.08 12.08 0.60 11.48 0.45 11.48 0.04 0.12 0.1 C17 0 0 0 0 0 0 0 11.39 11.39 0 11.39 0 11.39 0 0  C18 0 0 0 0 0 0 0 10.85 10.85 0 10.85 0 10.85 0 0  C19 0 0 0 0 0 0 0 10.31 10.31 0 10.31 0 10.31 0 0  C20 0 0 0 0 0 0 0 9.76 9.76 0 9.76 0 9.76 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
152 

Table E-14. Syngas Conditioning Area Stream Data for LT Scenario 

HT A400COND

CD35SGASCD36SGASCD37SGASCD38SGASCD39SGASCD40SGASCD41SGASCD42SGASCD44SGASCD45SGASCD80STMCD81STMCD99C

Temperature (C) 78 150 150 870 870 300 300 474 363 200 300 870 Pressure (bar) 26.50 26.50 26.50 26.50 25.81 25.81 25.81 25.12 25.12 25.12 25.00 27.00 Vapor Fraction 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 Volume Flow** (m3/sec) 2.18 2.64 2.64 7.08 10.44 5.24 1.83 2.46 5.98 4.44 1.15 2.26 

Mole Flow** (kmol/hr) 7066 7066 7063 7063 10153 10153 3554 3554 10153 10153 2313 2313 Mass Flow (tonnes/day) 2706 2706 2705 2705 3705 3705 1297 1297 3705 3705 1000 1000  H2O 30.18 30.18 30.18 30.18 850.65 850.65 297.73 118.99 671.91 671.91 1000 1000  CO 1575 1575 1575 1575 1854 1854 648.81 370.91 1576 157600 H2 168.38 168.38 168.38 168.38 222.13 222.13 77.75 97.75 242.13 242.1300 CO2 167.80 167.80 167.80 167.80 167.80 167.80 58.73 495.37 604.44 604.4400 O2 0 0 0 0 0 0 0 0 0 0 0 0  N2 0 0 0 0 0 0 0 0 0 0 0 0  AR 411.50 411.50 411.50 411.50 411.50 411.50 144.02 144.02 411.50 411.5000 CH4 149.06 149.06 149.06 149.06 52.13 52.13 18.25 18.25 52.13 52.1300 C2H6 28.49 28.49 28.49 28.49 7.72 7.72 2.70 2.70 7.72 7.72 0 0  C2H4 36.05 36.05 36.05 36.05 0.39 0.39 0.14 0.14 0.39 0.39 0 0  C6H6 39.50 39.50 39.50 39.50 39.50 39.50 13.83 13.83 39.50 39.50 0 0  C3 54.18 54.18 54.18 54.18 54.18 54.18 18.96 18.96 54.18 54.18 0 0  C4 34.31 34.31 34.31 34.31 34.31 34.31 12.01 12.01 34.31 34.31 0 0  H2S 0.13 0.13 0.01 0.01 0.01 0.01 0 0 0.01 0.01 0 0  NH3 1.52 1.52 0.30 0.30 0.30 0.30 0.11 0.11 0.30 0.30 0 0  TAR 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 0 0 0 0 0 0 0 0 0 0 0 0  C5 2.21 2.21 2.21 2.21 2.21 2.21 0.77 0.77 2.21 2.21 0 0  C6 2.37 2.37 2.37 2.37 2.37 2.37 0.83 0.83 2.37 2.37 0 0  C7 1.11 1.11 1.11 1.11 1.11 1.11 0.39 0.39 1.11 1.11 0 0  C8 1.14 1.14 1.14 1.14 1.14 1.14 0.40 0.40 1.14 1.14 0 0  C9 1.08 1.08 1.08 1.08 1.08 1.08 0.38 0.38 1.08 1.08 0 0  C10 1.13 1.13 1.13 1.13 1.13 1.13 0.40 0.40 1.13 1.13 0 0  C11 0.50 0.50 0.50 0.50 0.50 0.50 0.18 0.18 0.50 0.50 0 0  C12 0.44 0.44 0.44 0.44 0.44 0.44 0.16 0.16 0.44 0.44 0 0  C13 0.34 0.34 0.34 0.34 0.34 0.34 0.12 0.12 0.34 0.34 0 0  C14 0.20 0.20 0.20 0.20 0.20 0.20 0.07 0.07 0.20 0.20 0 0  C15 0 0 0 0 0 0 0 0 0 0 0 0  C16 0.04 0.04 0.04 0.04 0.04 0.04 0.01 0.01 0.04 0.04 0 0  C17 0 0 0 0 0 0 0 0 0 0 0 0  C18 0 0 0 0 0 0 0 0 0 0 0 0  C19 0 0 0 0 0 0 0 0 0 0 0 0  C20 0 0 0 0 0 0 0 0 0 0 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
153 

Table E-15. Hydroprocessing, Power Generation, and Air Separation Areas Stream Data for LT Sce

HT A500

HY56FT HY65FGAS HY70GASO HY71GASOHY80DIES HY81DIES HY90HYD A600 421 602 615 680 CW2CWREC HPW1HPSTM A700 AIR-1A AIR-A AIR-L

Temperature (C) 35 50 50 50 50 50 60 32 1161 344 30 85 23 170 393 -170 32 20 

Pressure (bar) 22.89 22.89 22.89 22.89 22.89 22.89 1.01 22.89 1.00 1.00 1.01 7.91 0.02 7.91 173.38 6.20 1.01 6.30 

Vapor Fraction 0.00 1.00 0.00 0.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00 0.00 0.89 0.00 1.00 1.00 1.00 1.00 Volume Flow** (m3/sec) 0.01 0.02 0 0 0 0 0.53 0.12 94.88 40.84 17.42 0.01 442.16 0 0.03 1.06 23.17 3.58 

Mole Flow** (kmol/hr) 69.40 57.67 31.78 31.78 37.88 37.88 69.96 394.75 2863 2863 2522 1709 1709 1709 1709 3162 3328 3328 Mass Flow (tonnes/day) 330.42 40.83 87.12 87.12 205.86 205.86 3.38 167.90 1955 1955 1746 740.10 740.10 740.10 740.10 2198 2313 2313 

 H2O 0.09 0 0 0 0 0 0 0 189.23 189.23 0 740.10 740.10 740.10 740.10 0 0 0  CO 0 0 0 0 0 0 0 64.47 0 0 0 0 0 0 0 0 0 0  H2 0 0 0 0 0 0 3.38 9.79 0 0 0 0 0 0 0 0 0 0  CO2 0 0 0 0 0 0 0 45.33 318.35 318.35 0 0 0 0 0 0 0 0  O2 0 0 0 0 0 0 0 0 80.99 80.99 410.78 0 0 0 0 508.87 535.65 535.65  N2 0 0 0 0 0 0 0 0 1336 1336 1336 0 0 0 0 1661 1748 1748  AR 0 0 0 0 0 0 0 30.86 30.86 30.86 0 0 0 0 0 28.22 29.71 29.71  CH4 0 11.56 0 0 0 0 0 4.22 0 0 0 0 0 0 0 0 0 0  C2H6 0 0 0 0 0 0 0 1.10 0 0 0 0 0 0 0 0 0 0  C2H4 0 0 0 0 0 0 0 0.03 0 0 0 0 0 0 0 0 0 0  C6H6 0 0 0 0 0 0 0 2.96 0 0 0 0 0 0 0 0 0 0  C3 0 29.27 0 0 0 0 0 4.75 0 0 0 0 0 0 0 0 0 0  C4 0 0 0 0 0 0 0 3.38 0 0 0 0 0 0 0 0 0 0  H2S 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NH3 0 0 0 0 0 0 0 0.02 0 0 0 0 0 0 0 0 0 0  TAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SULFUR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CARBON 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  STEAM 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  NO2 0 0 0 0 0 0 0 0 0.06 0.06 0 0 0 0 0 0 0 0  MEA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  AIR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  WAXES 131.46 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C5 11.92 0 0 0 0 0 0 0.18 0 0 0 0 0 0 0 0 0 0  C6 12.82 0 0 0 0 0 0 0.19 0 0 0 0 0 0 0 0 0 0  C7 13.53 0 0 0 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0  C8 13.88 0 87.12 87.12 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0  C9 13.95 0 0 0 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0  C10 13.98 0 0 0 0 0 0 0.09 0 0 0 0 0 0 0 0 0 0  C11 13.86 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C12 13.54 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C13 13.10 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C14 12.57 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C15 11.93 0 0 0 0 0 0 0.04 0 0 0 0 0 0 0 0 0 0  C16 11.48 0 0 0 205.86 205.86 0 0.04 0 0 0 0 0 0 0 0 0 0  C17 11.39 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C18 10.85 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C19 10.31 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  C20 9.76 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  BIOMASS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  ASH 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0  CHAR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 **Volumetric and Mole flow values do not include biomass, ash, or char
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November 2010

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Technical Report

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4. TITLE AND SUBTITLE

Techno-Economic Analysis of Biofuels Production Based on 

Gasification

5a. CONTRACT NUMBER

DE-AC36-08-GO28308

5b. GRANT NUMBER

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6. AUTHOR(S)

R.M. Swanson, J.A. Satrio, R.C. Brown, A. Platon, and D.D. Hsu

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NREL/TP-6A20-46587

5e. TASK NUMBER

BB07.7510

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National Renewable Energy Laboratory

1617 Cole Blvd.

Golden, CO 80401-3393

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NREL/TP-6A20-46587

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NREL

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14. ABSTRACT (Maximum 200 Words)

This study compares capital and production costs of two biomass-to-liquid production plants based on gasification. 

The first biorefinery scenario is an oxygen-fed, low-temperature (870°C), non-slagging, fluidized bed gasifier. The 

second scenario is an oxygen-fed, high-temperature (1,300°C), slagging, entrained flow gasifier. Both are followed by 

catalytic Fischer-Tropsch synthesis and hydroprocessing to naphtha-range (gasoline blend stock) and distillate-range 

(diesel blend stock) liquid fractions. Process modeling software (Aspen Plus) is utilized to organize the mass and 

energy streams and cost estimation software is used to generate equipment costs. Economic analysis is performed 

to estimate the capital investment and operating costs. Results show that the total capital investment required for nth 

plant scenarios is $610 million and $500 million for high-temperature and low-temperature scenarios, respectively. 

Product value (PV) for the high-temperature and low-temperature scenarios is estimated to be $4.30 and $4.80 per 

gallon of gasoline equivalent (GGE), respectively, based on a feedstock cost of $75 per dry short ton. Sensitivity 

analysis is also performed on process and economic parameters. This analysis shows that total capital investment 

and feedstock cost are among the most influential parameters affecting the PV. 

15. SUBJECT TERMS

biomass; biofuels; gasification; Fischer-Tropsch; techno-economic analysis; biomass-to-liquid; thermochemical

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