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

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

Contract No. DE-AC36-08GO28308

Process Design and Economics 

for Conversion of Lignocellulosic 

Biomass to Ethanol 

Thermochemical Pathway by Indirect 

Gasification and Mixed Alcohol 

Synthesis

A. Dutta, M. Talmadge, and J. Hensley

National Renewable Energy Laboratory

Golden, Colorado

M. Worley and D. Dudgeon

Harris Group Inc.

Atlanta, Georgia and Seattle, Washington

D. Barton, P. Groenendijk, D. Ferrari, and 

B. Stears

The Dow Chemical Company

Midland, Michigan

E.M. Searcy, C.T. Wright, and J.R. Hess 

Idaho National Laboratory

Idaho Falls, Idaho

Technical Report

NREL/TP-5100-51400

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

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

National Renewable Energy Laboratory

1617 Cole Boulevard

Golden, Colorado 80401

303-275-3000 • www.nrel.gov

Contract No. DE-AC36-08GO28308

Process Design and Economics 

for Conversion of Lignocellulosic 

Biomass to Ethanol 

Thermochemical Pathway by Indirect 

Gasification and Mixed Alcohol 

Synthesis

A. Dutta, M. Talmadge, and J. Hensley

National Renewable Energy Laboratory

Golden, Colorado

M. Worley and D. Dudgeon

Harris Group Inc.

Atlanta, Georgia and Seattle, Washington

D. Barton, P. Groenendijk, D. Ferrari, and 

B. Stears

The Dow Chemical Company

Midland, Michigan

E.M. Searcy, C.T. Wright, and J.R. Hess 

Idaho National Laboratory

Idaho Falls, Idaho

Prepared under Task No. BB07.3710

Technical Report

NREL/TP-5100-51400

May 2011
NOTICE

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Executive Summary 

The Energy Independence and Security Act (EISA) of 2007 established a mandate for 36 billion 

gallons of U.S. biofuel production per year by 2022. To meet this goal, the U.S. Department of 

Energy (DOE) promotes the development of ethanol and other liquid fuels from lignocellulosic 

feedstocks and funds fundamental and applied research to advance the state of biomass 

conversion technology. As a participant in the program, the National Renewable Energy 

Laboratory (NREL) investigates process improvements for cellulosic ethanol production and 

develops information that can be used to assess cellulosic ethanol’s cost-competitiveness and 

market penetration potential. 

This design report describes an up-to-date benchmark thermochemical conversion process that 

incorporates the latest research from NREL and other sources. It also includes improvements that 

are deemed experimentally achievable by 2012. This benchmark helps to quantify the impacts of 

core research at NREL and elsewhere and evaluates progress toward achieving technical and cost 

targets. The report will assist DOE in understanding the research areas that hold the greatest 

potential to positively impact the cost of ethanol production via gasification. 

Building on a design report published in 2007, NREL and its subcontractor Harris Group Inc. 

performed a complete review of the process design and economic model for a biomass-toethanol process via indirect gasification. Specifically, the process design involves conversion of 

woody biomass to ethanol, where biomass is converted to syngas via indirect steam gasification, 

and the syngas is cleaned, conditioned, and converted to mixed alcohols over a solid catalyst. 

Ancillary areas such as feed handling and drying, alcohol separation, steam and power 

generation, cooling water, and other utilities are also included in the design. 

The following major updates are included in the report: 

•	 Improved feedstock cost estimates derived from Idaho National Laboratory’s uniformformat feedstock handling system. 

•	 Current understanding of the conversion of tars and other hydrocarbons to carbon 

monoxide (CO) and hydrogen (H2) using biomass-derived syngas and steam reforming. 

•	 Inclusion of The Dow Chemical Company’s fundamental kinetic model for alcohol 

synthesis, which predicts the behavior of their proprietary alcohol synthesis catalyst. 

•	 Inclusion of a physical solvent for the removal of CO2 at high pressures. 

The conceptual design presented herein considers the economics of ethanol production, assuming 

the achievement of internal research targets for 2012 and nth-plant costs and financing. The 

design features a processing capacity of 2,205 U.S. tons (2,000 metric tonnes) of dry biomass per 

day and an ethanol yield of 83.8 gallons per dry U.S. ton of feedstock. The ethanol selling price 

corresponding to this design is $2.05 per gallon in 2007 dollars, assuming a 30-year plant life 

and 40% equity financing with a 10% internal rate of return and the remaining 60% debt 

financed at 8% interest. This ethanol selling price corresponds to a gasoline equivalent price of 

$3.11 per gallon based on the relative volumetric energy contents of ethanol and gasoline. A 

summary of the analysis is presented in the following table. 

iii 
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Process Engineering Analysis for Ethanol from Mixed Alcohol Synthesis 

2012 Case Using Technical Targets for Tar Reforming & Mixed Alcohol Synthesis
 

2,000 Dry Metric Tonnes Biomass per Day
 

Indirect Gasifier, Tar Reformer, Sulfur Removal, MoS2 Catalyst, Fuel Purification, Steam-Power Cycle
 

All Values in 2007$
 

Minimum Ethanol Selling Price (MESP) $2.05 per Gallon 

Contributions: Feedstock Costs $0.735 per Gallon 

Operating Costs & Credits $0.244 per Gallon 

Capital Charges & Taxes $1.069 per Gallon 

Gasoline-Equivalent Price $3.11 per Gallon 

EtOH Production at Operating Capacity 64.7 MM Gallons per Year 

EtOH Product Yield 83.8 Gallons per Dry US Ton Feedstock 

Mixed Alcohols Production at Operating Capacity 72.5 MM Gallons per Year 

Mixed Alcohols Product Yield 93.9 Gallons per Dry US Ton Feedstock 

Delivered Feedstock Cost $61.57 per Dry US Ton (Includes Capital Up to Throat of Gasifier) 

Internal Rate of Return (After-Tax) 10.0% 

Equity Percent of Total Investment 40.0% 

Capital Costs Operating Costs (¢ / Gallon Product)

 Gasification $43,250,000 Feedstock 73.5

 Tar Reforming & Quench $26,940,000 Natural Gas 0.0

 Acid Gas & Sulfur Removal $28,490,000 Catalysts 9.9

 Syngas Compression & Expansion $80,630,000 Olivine & Magnesium Oxide 0.7

 Alcohol Synthesis Reaction $41,420,000 Other Raw Materials 0.9

 Alcohol Separation $20,310,000 Waste Disposal 0.8

 Steam System & Power Generation $45,840,000 Electricity 0.0

 Cooling Water & Other Utilities $9,560,000 Fixed Costs 35.9 

Total Installed Equipment Cost (TIC) $296,450,000 Co-Product Credits -23.8 

Capital Depreciation 37.8 

Land (115 acres at $14000 per acre) 1,600,000 Average Income Tax 11.5 

Site Development 9,640,000 Average Return on Investment 57.5 

(% of ISBL) 4.0% 

Indirect Costs & Project Contingency 183,650,000 

(% of TIC) 62.0% Operating Costs ($ / Year) 

Feedstock $47,560,000 

Fixed Capital Investment (FCI) 491,350,000 Natural Gas $0 

Working Capital 24,490,000 Catalysts $6,380,000 

Total Capital Investment (TCI) 515,840,000 Olivine & Magnesium Oxide $0 

Other Raw Materials $380,000 

Total Installed Equipment Cost per Annual Gallon Waste Disposal $520,000 

of Ethanol Product 4.58 Electricity $0 

of Mixed Alcohol Product 4.09 Fixed Costs $23,240,000 

Co-Product Credits at $1.88 per Gallon -$15,380,000 

Fixed Capital Investment per Annual Gallon Capital Depreciation $24,490,000 

of Ethanol Product 7.59 Average Income Tax $7,450,000 

of Mixed Alcohol Product 6.78 Average Return on Investment $37,210,000 

Loan Interest Rate 8.0% Total Plant Electricity Usage (KW) 64,356 

Loan Term (Years) 10 Electricity Produced Onsite (KW) 64,411 

Electricity Purchased from Grid (KW) 0 

Maximum Yields Based on Feedstock Carbon Content Electricity Sold to Grid (KW) 55 

Theoretical Ethanol Production (MM Gal / Year) 159.0 

Theoretical Ethanol Yield (Gal / Dry Ton) 205.9 Steam Plant + Turboexpander Power Generated (hp) 86,377 

Current Ethanol Yield (Actual / Theoretical) 40.7% Used for Main Compressors (hp) 63,374 

Other Electical Consumption (hp) 23,003 

Gasifier Efficiency - HHV % 74.3 

Gasifier Efficiency - LHV % 73.9 Plant Electricity Use (KWh / Gal EtOH) 8.36 

Overall Plant Efficiency - HHV % 46.5 Gasification & Reforming Steam Use (lb / Gal EtOH) 8.68 

Overall Plant Efficiency - LHV % 44.9 

Specific Operating Conditions 

Plant Operating Hours per Year 8410 Feed Rate Dry Tonnes / Day 2,000 

On-Stream Percentage 96.0% Dry Tons / Day 2,205 

Feedstock Cost $ / Dry Ton $61.57 

Excel File: R236H-V29.xls $ / Moisture & Ash Free Ton $62.14 

iv 
 

 

 

 

Acknowledgements 

We thank the reviewers and other contributors to this report. Reviewers and other contributors 

are listed below by company/organization name. Reviewer comments on the draft version of the 

report are presented in Appendix O along with NREL authors’ responses. 

List of reviewers 

Black & Veatch Jon Erickson 

British Petroleum Nicholas Petrellis (retiree with 30 years experience) 

Colorado School of Mines Rob Braun 

ConocoPhillips Alex Platon 

Coskata, Inc. Richard Tobey 

Goldman Sachs Brian Bolster 

National Renewable Energy Laboratory Andy Aden 

Richard Bain 

Mary Biddy 

Ryan Davis 

Kristiina Iisa 

Kim Magrini-Bair 

Steven Phillips 

Joan Tarud 

Pacific Northwest National Laboratory Susanne Jones 

Iva Jovanovic 

Yunhua Zhu 

Rentech, Inc. George Apanel 

URS Corporation John Jechura 

Weyerhaeuser John Tao 

Editor 

National Renewable Energy Laboratory Sara Havig 

Process flow diagram development 

Harris Group Inc. Diane Hartley 

Susanne Korinek 

Ray Lanier 

Hung Tran 

Joan Walker 

We thank the U.S. Department of Energy for funding and supporting this work. 

v 
 

Table of Contents 

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

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

1.2 Technoeconomic Analysis Approach ..................................................................................3
 

1.3 Definition of nth-Plant Economics .......................................................................................4
 

1.4 Estimation of Capital Costs .................................................................................................6
 

2 Plant Design Basis ............................................................................................................................... 9
 

2.1 Feedstock and Plant Size .....................................................................................................9
 

2.2 Process Overview...............................................................................................................11
 

2.3 Aspen Plus Model..............................................................................................................14
 

3 Process Design and Cost Estimation............................................................................................... 15
 

3.1 Area 100: Feed Handling and Drying................................................................................15
 

3.1.1 Area 100 Overview ...................................................................................................15
 

3.1.2 Area 100 Design Basis..............................................................................................17
 

3.1.3 Area 100 Equipment Cost Estimations .....................................................................17
 

3.2 Area 200: Gasification .......................................................................................................18
 

3.2.1 Area 200 Overview ...................................................................................................19
 

3.2.2 Area 200 Design Basis..............................................................................................21
 

3.2.3 Area 200 Equipment Cost Estimations .....................................................................21
 

3.3 Area 300: Gas Cleanup ......................................................................................................23
 

3.3.1 Area 300 Overview ...................................................................................................23
 

3.3.2 Area 300 Design Basis..............................................................................................25
 

3.3.3 Area 300 Equipment Cost Estimations .....................................................................26
 

3.4 Area 400: Alcohol Synthesis, Syngas Compression and Acid Gas Processing.................26
 

3.4.1 Area 400 Overview ...................................................................................................26
 

3.4.2 Area 400 Design Basis..............................................................................................33
 

3.4.3 Area 400 Equipment Cost Estimations .....................................................................37
 

3.5 Area 500: Alcohol Separation............................................................................................37
 

3.5.1 Area 500 Overview ...................................................................................................38
 

3.5.2 Area 500 Design Basis..............................................................................................40
 

3.5.3 Area 500 Equipment Cost Estimations .....................................................................41
 

3.6 Area 600: Steam System and Power Generation ...............................................................41
 

3.6.1 Area 600 Overview ...................................................................................................41
 

3.6.2 Area 600 Design Basis..............................................................................................42
 

3.6.3 Area 600 Equipment Cost Estimations .....................................................................43
 

3.7 Area 700: Cooling Water and Other Utilities ....................................................................43
 

3.7.1 Area 700 Overview ...................................................................................................43
 

3.7.2 Area 700 Design Basis..............................................................................................44
 

3.7.3 Area 700 Cost Estimation .........................................................................................44
 

3.8 Additional Process Design Information.............................................................................45
 

3.9 Pinch Analysis ...................................................................................................................45
 

3.10 Energy Balance ................................................................................................................46
 

3.11 Sustainability Metrics ......................................................................................................48
 

4 Process Economics ........................................................................................................................... 51
 

4.1 Total Capital Investment (TCI)..........................................................................................51
 

4.2 Variable Operating Costs...................................................................................................53
 

vi 
 vii 

4.3 Fixed Operating Costs........................................................................................................54

4.4 Value of Higher Alcohol Co-Products ...............................................................................55

4.5 Minimum Ethanol Selling Price.........................................................................................55

5 Process Economics and Sensitivity Analyses ................................................................................ 56

5.1 Financial Scenarios ............................................................................................................58

5.2 Feedstock ...........................................................................................................................58

5.3 Gasification and Gas Cleanup ............................................................................................58

5.4 Alcohol Synthesis ..............................................................................................................58

5.5 Co-Product Value...............................................................................................................59

6 Conclusions ........................................................................................................................................ 59

7 Future Work ......................................................................................................................................... 59

8 References .......................................................................................................................................... 60

Appendix A. Equipment Description and Design Basis Summary ...................................................... 64

Appendix B. Individual Equipment Cost Summary ............................................................................... 74

Appendix C. Discounted Cash Flow Rate of Return (DCFROR) and Operating Cost Summary ...... 83

Appendix D. Process Parameters and Operation Summary ................................................................ 90

Appendix E. Process Flow Diagrams (PFDs) and Material and Energy Balances ............................. 93

Appendix F. Feedstock Supply Overview from Idaho National Laboratory (INL) ............................ 121

Appendix G. Synthesis Gas and Char Correlations for Indirect Gasifier .......................................... 128

Appendix H. Equipment Lists and Cost Data for Package Systems ................................................. 132

Appendix I. Technology and Cost Targets for Gasification, Tar Reforming, and Alcohol 

Synthesis ........................................................................................................................................... 137

Appendix J. Fuel-Grade Ethanol Specifications .................................................................................. 141

Appendix K. Direct Costs Included in Equipment Installation Factors from Harris Group Inc. ..... 143

Appendix L. Plant Footprint and Land Cost Calculation from Harris Group Inc. ............................. 149

Appendix M. Pinch Analysis and Heat Integration System Design ................................................... 151

Appendix N. Summary of Design Report Assumptions ...................................................................... 154

Appendix O. Reviewer Comments on Draft Design Report and Responses .................................... 157
 

 

 

 

List of Figures 

Figure 1. Approach to thermochemical ethanol process design and analysis................................. 3
 

Figure 2. Chemical Engineering’s Plant Cost Index data............................................................... 7
 

Figure 3. Current and predicted availability of biomass................................................................. 9
 

Figure 4. Simplified flow diagram for the thermochemical process ............................................ 11
 

Figure 5. Woody biomass feedstock supply system ..................................................................... 16
 

Figure 6. Simplified process flow diagram of gasification area ................................................... 19
 

Figure 7. Simplified process flow diagram of gas cleanup area ................................................... 23
 

Figure 8. Simplified process flow diagram of alcohol synthesis area .......................................... 27
 

Figure 9. The Dow Chemical Company model predictions compared to experimental data ....... 29
 

Figure 10. Simplified flow diagram of the H2S/CO2 removal and sulfur recovery...................... 31
 

Figure 11. Maximum operating pressure for syngas turbo-expander........................................... 32
 

Figure 12. Simplified process flow diagram of alcohol separation area....................................... 38
 

Figure 13. Pinch analysis composite curve................................................................................... 46
 

Figure 14. Overall energy analysis (dry wood LHV basis) .......................................................... 47
 

Figure 15. Cost contribution details from each process area ........................................................ 56
 

Figure 16. Results of sensitivity analyses..................................................................................... 57
 

List of Tables 

Table 1. Summary of nth-Plant Assumptions for Technoeconomic Analysis................................. 5
 

Table 2. Comparison of Market Studies and TE Analyses for Existing Biofuels .......................... 5
 

Table 3. Direct Cost Factors for Estimating Installed Equipment Costs [16] ................................ 8
 

Table 4. Ultimate Analysis of Woody Biomass Feedstock .......................................................... 10
 

Table 5. Design Basis for Feed Handling and Drying Area ......................................................... 17
 

Table 6. Unit Cost Breakdown for Woody Biomass Feedstock (Southern Pine)......................... 18
 

Table 7. Design Basis for Gasification Area ................................................................................ 21
 

Table 8. Estimates for Gasification and Tar Reforming Facilities ............................................... 22
 

Table 9. Cost Estimate for Gasification Area ............................................................................... 22
 

Table 10. Target Performance of Reforming Catalyst.................................................................. 25
 

Table 11. Reformer Conditions, Gas Compositions, and Reactor Basis for the Base Case ......... 25
 

Table 12. Cost Estimate for Gas Cleanup Area ............................................................................ 26
 

Table 13. Target and Design Basis Parameters for Alcohol Synthesis Reactor ........................... 34
 

Table 14. Design Parameters for Compression, Power Recovery, and H2S/CO2 Systems .......... 35
 

Table 15. Sources of Equipment Costs for Alcohol Synthesis Area ............................................ 37
 

Table 16. Cost Estimate for Alcohol Synthesis Area ................................................................... 37
 

Table 17. Design Basis Parameters for Alcohol Separation Area ................................................ 40
 

Table 18. Cost Estimate for Alcohol Separation Area.................................................................. 41
 

Table 19. Plant Power Requirements............................................................................................ 42
 

Table 20. Cost Estimate for Steam System and Power Generation Area ..................................... 43
 

Table 21. Design Basis for Cooling Water System ...................................................................... 44
 

Table 22. Cost Estimate for Cooling Water and Other Utilities Area .......................................... 44
 

Table 23. Utility and Miscellaneous Design Information............................................................. 45
 

Table 25. Comparison of Water Usage for Transportation Fuel Processes.................................. 50
 

Table 26. General Cost Factors in Determining Total Installed Equipment Costs....................... 51
 

viii 
Table 27. Cost Factors for Indirect Costs ..................................................................................... 52
 

Table 28. Project Cost Worksheet ................................................................................................ 52
 

Table 29. Summary of Variable Operating Costs......................................................................... 53
 

Table 30. Salary Costs for Plant Employees................................................................................. 54
 

Table 31. Other Fixed Operating Costs ........................................................................................ 54
 

Table 32. Higher Alcohols Co-Product Pricing............................................................................ 55
 

ix 
 

 

 

 

 

 

 

 

 

 

List of Acronyms
 

ASME American Society of Mechanical ISBL inside battery limits 

Engineers kWh kilowatt-hour 

ASTM American Society of Testing LHV lower heating value 

and Materials MAF moisture and ash free 

bbl barrel = 42 gallons MA mixed alcohols 

BCL Battelle Columbus Laboratory MAS mixed alcohol synthesis 

BFW boiler feed water MASP minimum alcohols selling price 

BPD barrels per day MeOH methanol 

Btu British thermal unit MESP minimum ethanol selling price 

CFM cubic feet per minute MoS2 molybdenum disulfide 

CH4 methane MgO magnesium oxide 

CO carbon monoxide MTBE methyl-tertiary-butyl ether 

Co cobalt MTPD metric tonnes per day 

CO2 carbon dioxide MW megawatts 

DCFROR discounted cash flow rate of return NREL National Renewable Energy Laboratory 

DEPG dimethyl ether of polyethylene glycol NRTL non-random two liquid activity coefficient 

DOE U.S. Department of Energy method 

EDI electrodeionization OBP DOE’s Office of Biomass Program 

EIA Energy Information Administration PFD process flow diagram 

EISA Energy Independence and Security PNNL Pacific Northwest National Laboratory 

Act PPMV parts per million by volume 

EtOH ethanol PPMW parts per million by weight 

FCC fluid catalytic cracking POX partial oxidation 

FCI fixed capital investment psia pounds per square inch (absolute) 

FT Fischer-Tropsch psig pounds per square inch (gauge) 

FY fiscal year RKS-BM Redlich-Kwong-Soave equation of state 

GHSV gas hourly space velocity with Boston-Mathias modifications 

GJ gigajoule SMR steam methane reformer 

gpm gallons per minute SOT State of Technology 

H2 hydrogen TIC total installed cost 

H2S hydrogen sulfide TDC total direct cost 

HAS higher alcohol synthesis TPD short (U.S.) tons per day 

HHV higher heating value TPEC total purchased equipment cost 

hp horsepower WGS water gas shift 

IRR internal rate of return WWT wastewater treatment 

INL Idaho National Laboratory 

IPE Icarus Process Evaluator 

(software from Aspen) 

x 
 

 

 

 

 

 

 

 

 

1 Introduction 

1.1 Motivation and Background 

The Energy Independence and Security Act of 2007 (EISA) [1] established a mandate for the 

annual U.S. production of 36 billion gallons of renewable fuel by 2022, of which 15 billion 

gallons may come from traditional or corn-based technologies and at least 16 billion gallons 

must be derived from lignocellulosic biomass sources. In order to achieve this goal the U.S. 

Department of Energy (DOE) has promoted the development of ethanol and other liquid fuels 

from lignocellulosic feedstocks by sponsoring programs in fundamental and applied research that 

aim to advance biomass conversion technologies. 

One of DOE’s research platforms is thermochemical conversion of lignocellulosic biomass to 

ethanol via gasification. DOE has set a target for making available technologies that, when 

integrated, will enable the production of ethanol that is cost-competitive with gasoline on an 

energy equivalent basis, without subsidies or policy initiatives, by the year 2012. To achieve this 

goal, DOE is sponsoring research in areas where commercial technology is not currently 

available, such as cleanup of biomass-derived syngas and synthesis of ethanol from syngas. This 

research is conducted at national laboratories, universities, and private companies in conjunction 

with engineering companies. 

In order to assess the economic viability and potential of the conversion process, the National 

Renewable Energy Laboratory (NREL) and its partners have developed a complete process 

model with economic projections to estimate an absolute plant-gate price for ethanol. This plantgate price is referred to as the minimum ethanol selling price (MESP). The MESP can be used by 

policymakers and DOE to assess the cost-competitiveness and market penetration potential of 

cellulosic ethanol in comparison with petroleum-derived fuels and corn-based ethanol. 

The process design and economic analysis also directs biomass conversion research by 

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

research and anticipated results can be translated into a new MESP value that can be compared to 

the benchmark value documented in this report. This process helps to quantify the impact of core 

research at NREL and elsewhere and to track progress toward meeting competitive cost targets. 

It also allows DOE to make decisions about which research proposals have the greatest potential 

to reduce MESP. 

This report builds upon a previous design report released in 2007 [2], written by engineers from 

NREL and Neoterics International, Inc. For the present report, NREL contracted Harris Group 

Inc. to provide engineering support through estimates and reviews of the equipment and raw 

material cost assumptions used in the process design and economic analysis. This report reflects 

NREL’s latest envisioned thermochemical ethanol process design and includes the latest research 

and development progress in three key conversion areas (gasification, gas cleanup, and alcohol 

synthesis) and improvements in other areas such as heat integration and utilities. 

Gasification involves the devolatilization and chemical conversion of biomass to a mediumcalorific-value gas in an atmosphere of steam and/or oxygen. There are two primary classes of 

gasifiers – partial oxidation or direct gasifiers, and indirect steam blown gasifiers. Partial 

oxidation (POX) gasifiers, also called directly-heated gasifiers, use heat from the reaction 

1
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

between oxygen and organic compounds to provide the necessary energy to devolatilize biomass 

and to react with carbon-rich chars to break their complex molecules into gases. POX gasifiers 

require high-purity oxygen provided by air separation units, which require significant capital 

investment. The capital intensity of these systems makes direct gasification processes more costprohibitive for smaller scale biomass conversion plants [3]. 

Indirect steam gasifiers are heated via heat transfer from a hot solid or through a heat transfer 

surface. By-product char and portions of the product gas can be combusted with air (external to 

the gasifier itself) to provide the energy for gasification. Although steam gasifiers do not require 

an air separation unit, gas compression is typically required for downstream unit operations 

because most indirect gasifiers operate at low pressure. 

Indirect steam gasification was chosen as the basis for the gasifier design because 

technoeconomic studies have shown that it has an economic advantage over oxygen-blown hightemperature slagging [4] or lower temperature dry-ash direct gasification [5] processes. Similar 

conclusions were drawn for the production of methanol and hydrogen from biomass [6]. 

A constraint was incorporated into this design requiring the process to remain “electrical energy 

neutral.” That is, the process was designed without allowing for purchase of coal, natural gas, or 

electricity for plant operations. Because the heat and power requirements of the process cannot 

be met through char combustion alone, the following options were considered: 

•	 Additional biomass could be used to heat and power the process. This option increases 

the biomass processing rate beyond 2,000 tonnes per day and potentially voids some of 

the feedstock cost assumptions. 

•	 Fossil fuels (coal or natural gas) could be added directly to provide additional fuel. 

•	 Some syngas could be diverted from liquid fuel production for heat and power 

production. This option makes the design energy self-sufficient but lowers the overall 

product yield. 

It was decided that the first and second options would not be considered and that the third option 

would be used to satisfy balance-of-plant heat and power requirements. In designs that result in 

excess heat or power production, energy in the form of electricity will be exported to the grid for 

a co-product credit. 

NREL worked with Harris Group Inc. to identify realistic process configurations and equipment 

costs for critical, high-priced equipment in each process area. Aspen Icarus Process Evaluator 

(IPE) 2006.5 was utilized for cost estimations for generic, lower-cost equipment like knock-out 

drums and standard heat exchangers. 

The biomass conversion parameters used in this design are based on a slate of research targets 

that NREL and DOE have specified for the period spanning 2006 to 2012. These targets are 

discussed in this report and summarized in Appendix I. The economics of the process design use 

the best available equipment and raw material cost estimates with an “nth-plant” project cost 

structure and financing. The 2012 nth-plant MESP computed in this report is $2.05 per gallon in 

2
 
 

 

 

 

 

 

 

2007 dollars. This MESP value corresponds to a gasoline equivalent price of $3.11 per gallon. 

The gasoline equivalent price is determined by multiplying the MESP by the ratio of volumetric 

energy contents of gasoline and ethanol ($2.05 MESP * 115,500 Btu per gallon of gasoline / 

76,000 Btu per gallon of ethanol). 

Future modifications to the process design presented here will be reflected annually through 

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

associated costs incorporate the most current data from NREL, other DOE-funded research 

programs, and external sources, including up-to-date equipment and installation costs. 

1.2 Technoeconomic Analysis Approach 

The general approach used in the process design, process model, and economic analysis is 

depicted in Figure 1. This design report includes information from completed and ongoing 

research at NREL and other national laboratories, detailed reviews of commercially available 

technologies, process modeling using Aspen Plus software, equipment cost estimation through 

vendor quotes and Aspen Icarus software, and discounted cash flow analysis. The 

technoeconomic analysis ultimately provides a minimum ethanol selling price (MESP) based on 

the financial assumptions in Table 1. 

Engineering 

Consultation on Process 

Configuration 

Process Flow Diagrams 

Results from 

DOE / NREL Sponsored 

Research Programs 

Rigorous Process 

Simulations 

(Aspen Plus) 

Material and Energy 

Balances 

Capital and Operating 

Cost Estimates 

Discounted Cash 

Flow Analysis 

Project Financing 

Assumptions 

Minimum Ethanol 

Selling Price (MESP) 

Third-Party Engineering 

Studies on Commercial 

Processes 

Process Estimates 

from Commercial 

Technology Licensors 

Vendor Cost 

Quotations 

Cost Estimates from 

Engineering Firm 

(Harris Group Inc) 

Results from Cost 

Estimating Software 

(Aspen Icarus) 

Figure 1. Approach to thermochemical ethanol process design and analysis 

3 
 

 

 

 

 

 

 

 

 

 

 

 

Aspen Plus Version 2004.1 was used to develop mass and energy balances for the process. The 

plant operations are separated into seven major process areas: 

• Area 100: Feed handling and drying 

• Area 200: Gasification 

• Area 300: Synthesis gas cleanup 

• Area 400: Alcohol synthesis, syngas compression, and acid gas removal 

• Area 500: Alcohol separation 

• Area 600: Steam system and power generation 

• Area 700: Cooling water and other utilities 

1.3 Definition of nth-Plant Economics 

The technoeconomic analysis reported here uses nth-plant economics. The key theoretical 

assumption associated with nth-plant economics is that several plants using the same technology 

have already been built and are operating. In other words, the assumption reflects a future in 

which a successful industry has been established with many operating plants. Because the 

technoeconomic model is a tool used primarily for (1) studying new process technologies or (2) 

comparing integrated schemes in order to comment on their relative economic impact, it is 

prudent to ignore artificial inflation of project costs associated with risk financing, longer startups, equipment overdesign, and other costs associated with pioneer plants, as these overshadow 

the real economic impact of advances in conversion science or process engineering research. At 

the very least, nth-plant economics should help to provide justification and support for early 

technology adopters and pioneer plants. 

Because equipment costs in this design report have been estimated explicitly, the nth-plant 

assumptions apply primarily to the factored cost model used to determine the total capital 

investment from the purchased equipment cost and to the assumptions applied for plant 

financing. The nth-plant assumption also applies to operating parameters, such as process uptime 

and start-up time. A summary of the nth-plant assumptions applied in this report are listed in 

Table 1. These financial assumptions are consistent with assumptions used for other economic 

analyses done for DOE’s Office of Biomass Program (OBP), with some deviations to reflect the 

uniqueness of this process. 

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Table 1. Summary of nth-Plant Assumptions for Technoeconomic Analysis
 

Description of Assumption Assumed Value
 

Internal rate of return (IRR) 10% 

Plant financing by equity / debt 40% / 60% of total capital investment 

Plant life 30 years 

Income tax rate 35% 

Interest rate for debt financing 8.0% annually 

Term for debt financing 10 years 

Working capital cost 5.0% of fixed capital investment 

(excluding land purchase cost) 

Depreciation schedule 7-Year MACRS schedule [7] 

Construction period (spending schedule) 3 years (8% Y1, 60% Y2, 32% Y3) 

Plant salvage value No value 

Start-up time 3 months 

Revenue and costs during start-up Revenue = 50% of normal 

Variable costs = 75% of normal 

Fixed costs = 100% of normal 

On-stream percentage after start-up 96% (8,410 operating hours per year) 

Examples have shown that technoeconomic models can reasonably predict the production costs 

of biofuels. Tao and Aden [8] performed a survey of technoeconomic models of existing biofuels 

production (corn ethanol, sugarcane ethanol, and conventional soy biodiesel) from published 

literature. These studies were normalized to a consistent year-dollar value and feedstock cost 

(where applicable) and compared to published market studies of the same biofuels. This 

comparison, shown in Table 2, indicates that technoeconomic (TE) models were able to predict 

the actual cost of production of these biofuels within an accuracy of ±20%. The variability in 

accuracy is largely a result of the variability in factored capital cost estimates used in these 

models. 

Table 2. Comparison of Market Studies and TE Analyses for Existing Biofuels 

Fuel Market Study TE Model Error in TE 

Prediction 

Corn ethanol $1.53/gal [9] $1.54/gal [10] 0.7 % 

Sugarcane ethanol $1.14/gal [9] $1.29/gal [11, 12] 13.2% 

Soy biodiesel $2.15/gal [13] $2.55/gal [14] 18.6% 

The consistency of TE predictions compared to market study values is not surprising because the
 

TE model developers were able to compare their results to real economics and make adjustments 

5
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

as necessary. Also, the feedstock costs for the fuels presented in Table 2 make up a high 

percentage of the total production cost, making errors in the remaining production cost estimates 

less impactful on the fuel selling price. Developing a technoeconomic model for a precommercial technology, such as the woody biomass-to-ethanol process described in this report, 

is based on less specific information and has to be rooted in current understanding of the state of 

the technology, feasible improvements at the time of the analysis, and good engineering practice. 

There will be more uncertainties in the cost predictions from such studies. Using the nth-plant 

philosophy, it is assumed that the pre-commercial technoeconomic model reflects the production 

economics for the technology’s mature future. 

1.4 Estimation of Capital Costs 

Capital costs were estimated using a variety of resources. For sub-processes that utilize welldeveloped technologies and can be purchased as modular packages (e.g., sulfur recovery unit and 

refrigeration), an overall package cost was used instead of the sum of costs for individual pieces 

of equipment. Costs for common process equipment (e.g., tanks, drums, pumps, and simple heat 

exchangers) were estimated using Aspen IPE 2006.5 costing software, which uses the Q1 2007 

cost basis. Unit operations that are specific to thermochemical ethanol production (e.g., gasifier, 

tar reformer, and alcohol synthesis reactor) were priced through vendor quotations. It is 

important to note that capital cost estimates for pre-commercial technologies like the gasifier and 

tar reformer were derived from current vendor quotes for first-of-its-kind fabrications, which 

should be higher than costs for mature nth-plant technologies. The cost assumptions for the 

gasifier and tar reformer are discussed in detail in Section 3.2.3. 

The original (base) purchased equipment costs reflect the base case for equipment size and cost 

year. Equipment sizes required for the process may vary from the original base case, requiring 

adjustment of the equipment costs. Instead of re-pricing equipment after minor changes in size, 

exponential scaling is applied to adjust the purchased equipment costs using Equation 1: 

n 

Scale-Up Capacity (Eq. 1) Scale-Up Equipment Cost = Base Equipment Cost Base Capacity 

The characteristic scaling exponent, n, is typically in the range of 0.6 to 0.7 for process 

equipment. The sizing parameters are based on a characteristic of the equipment related to 

production capacity, such as inlet flow for a process vessel or heat transfer duty for a heat 

exchanger. Equation 1 assumes that all other process parameters (pressure, temperature, etc.) 

remain constant relative to the base case. Scaling exponents were determined from the following 

sources: 

•	 Vendors’ estimates of scaling exponent or inference from vendor quotes when multiple 

quotes were available for equipment of various processing capacities. 

•	 Development of correlations by multiple estimates from Aspen IPE software. 

•	 Standard reference from published sources such as Garrett [15], Peters, Timmerhaus and 

West [16], and Perry et al. [17]. 

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When cost data were not available in 2007 dollars, costs were adjusted with Chemical 

Engineering’s (CE) Plant Cost Index [18] using Equation 2: 

Corrected Equipment Cost = Base Equipment Cost (Eq. 2) 

2007 Cost Index Value 

Base Year Cost Index Value 

The CE indices used in this study are listed and plotted in Figure 2. The index data show a sharp 

increase after 2003 due to increases in global steel demand and a dip in 2009 due to the global 

recession. The August 2010 index was used for 2010 since that was roughly the time when most 

of the recent cost quotes were obtained. 

Year Index Year Index 

1990 357.6 2001 394.3 

1991 361.3 2002 395.6 

1992 358.2 2003 402.0 

1993 359.2 2004 444.2 

1994 368.1 2005 468.2 

1995 381.1 2006 499.6 

1996 381.7 2007 525.4 

1997 386.5 2008 575.4 

1998 389.5 2009 521.9 

1999 

2000 

390.6 

394.1 

2010 

(August) 549.5 350 

400 

450 

500 

550 

600 

1985 1990 1995 2000 2005 2010 2015 

Figure 2. Chemical Engineering’s Plant Cost Index data 

Once the total purchased equipment costs (TPEC) were determined, scaled, and time-corrected, 

an equipment installation factor was applied to estimate the total installed cost (TIC) for the 

equipment or process unit including associated piping, instrumentation and controls, electrical 

systems, buildings, yard improvements, and direct labor per Equation 3, where fInstallation is the 

installation factor: 

Total Installed Cost (TIC) = f Installation * Total Purchased Equipment Cost (TPEC) (Eq. 3) 

Where possible, specific installation factors from vendor quotes or published data were applied 

to equipment or process units to estimate the total installed cost (TIC). When specific data was 

not available for equipment or package units, a default installation factor of 2.47 was applied to 

determine TIC. This value is based on the direct cost factor method presented by Peters, 

Timmerhaus, and West [16] for a solid-fluid processing plant, which is summarized in Table 3. 

Details for TPEC, installation factors, and TIC values for each process area are presented in 

Section 3. Harris Group Inc. prepared a detailed list of direct costs included in the equipment 

installation factors. The list is provided in Appendix K of this report. 

7
 
 

 

 

Table 3. Direct Cost Factors for Estimating Installed Equipment Costs [16] 

Direct Cost Description Cost as Fraction of TPEC 

Total purchased equipment cost 1.00 

Purchased equipment installation 0.39 

Instrumentation and controls 0.26 

Piping 0.31 

Electrical systems 0.10 

Buildings 0.29 

Yard improvements 0.12 

Total installation factor (fInstallation) 2.47 

Once the scaled TICs were determined, overhead and contingency factors were applied to 

determine a total capital investment (TCI) cost. The TCI along with plant operating expenses 

serves as the basis for the discounted cash flow analysis. The discounted cash flow analysis then 

yields the MESP, the primary metric by which the thermochemical ethanol process can be 

compared to alternate fuel production designs. 

8
 
 

 

 

 

 

 

 

 

 

 

 

2 Plant Design Basis 

2.1 Feedstock and Plant Size 

Forest resources were chosen as the primary feedstock for the thermochemical ethanol process 

design based upon expected availability from a 2005 joint DOE-USDA report assessing the 

feasibility of a billion-ton annual supply of biomass as feedstock for the bioenergy and 

bioproducts industries [19]. The assessment describes the potential for biomass availability to 

increase significantly with advancements in technology such as higher crop yields and improved 

residue collection. Potential biomass production rates are provided in Figure 3. The data set 

“High Yield Growth With Energy Crops” represents availability projections based on 

improvements in agricultural technology and the creation of a new energy crop industry. The 

data set “High Yield Growth Without Energy Crops” considers only improvements in 

agricultural technology. The data set “Existing and Unexploited Resources” represents biomass 

availability as of 2005. 

Studies for biochemical ethanol production have focused on using agricultural resources such as 

corn stover as feedstocks due to their lower lignin content relative to forest resources. Lignin 

does not readily break down in biochemical processes. Therefore, feedstocks with higher lignin 

content result in lower ethanol yields in biochemical processes. Because the thermochemical 

process by gasification is able to convert all feedstock carbon, including lignin, high-lignin forest 

resources serve as the base feedstock for the thermochemical plant design. 

Figure 3. Current and predicted availability of biomass 

The ultimate analysis for the feed considered in this study is presented in Table 4. The 

information sources for Table 4 indicate marginal differences in the ultimate analysis between 

hybrid poplar, used in the 2007 report [2], and pine, used as the basis for Idaho National 

Laboratory’s (INL) feedstock model. While this design maintains the wood composition from 

the 2007 report, the sulfur content was decreased from 0.09% to 0.03%, which is more 

9
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

representative of woody material. The assumed moisture content of the biomass feedstock prior 

to drying is 35 wt % at the plant gate, with an additional 5 wt % moisture loss during in-plant 

processing, per specifications from INL. A detailed description of INL’s feedstock supply system 

is provided in Appendix F. The effects of changing compositions and moisture contents are also 

examined as part of the sensitivity analysis, presented in Section 5. 

Table 4. Ultimate Analysis of Woody Biomass Feedstock 

Component Weight % (Dry Basis [20, 21]) 

Carbon 50.94 

Hydrogen 6.04 

Nitrogen 0.17 

Sulfur 0.03 

Oxygen 41.90 

Ash 0.92 

Heating valuea (Btu/lb) 8,601 HHV 

7,996 LHV 

a Calculated using the Aspen Plus Boie correlation. 

The design capacity for this study is 2,000 dry metric tonnes per day (2,205 dry tons per day), 

which matches that of previous design reports for both biochemical [22] and thermochemical [2] 

ethanol processes. With an expected 8,410 operating hours per year (96% on-stream factor / 

availability), the annual feedstock requirement is approximately 700,000 dry metric tonnes per 

year (772,000 dry tons per year). The assumed on-stream factor allows approximately 15 days of 

planned and unplanned downtime per year, which is consistent with mature (nth-plant) petroleum 

refining technologies such as fluid catalytic cracking (FCC) units [23]. The impact of lower onstream factors is shown as part of the sensitivity studies in Section 5. The feedstock is assumed 

to be delivered to the plants from satellite storage and processing facilities, details of which are 

described in Appendix F. 

The delivered feedstock cost was estimated using INL’s model. This cost of $61.57 per dry ton 

includes all capital costs, operating costs, and dry matter losses associated with feed delivery, 

drying, and handling as determined by INL. Up to the throat of the gasifier, the only process 

input that is added to the feedstock is flue gas waste heat for drying the feedstock from 30% 

moisture (after in-plant handling) to 10% moisture (at the gasifier). 

Impacts of plant size and feedstock cost on the MESP are included in the sensitivity analysis in 

Section 5. 

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2.2 Process Overview 

A block flow diagram of the current design is shown in Figure 4. Detailed process flow diagrams 

(PFDs) with material and energy balances are provided in Appendix E. 

Figure 4. Simplified flow diagram for the thermochemical process 

The proposed thermochemical ethanol production process includes the following steps: 

• Feed handling and preparation. The biomass feedstock is dried to 10 wt % moisture 

using hot flue gases from the char combustor and tar reformer catalyst regenerator. 

• Gasification. Biomass is indirectly gasified. Heat for the gasification reactions is supplied 

by circulating synthetic olivine1 sand that is pre-heated in a char combustor and fed to the 

gasifier. Conveyors and hoppers feed biomass to the low-pressure entrained flow gasifier. 

Steam is injected into the gasifier to stabilize the flow of biomass and olivine through the 

gasifier. Within the gasifier, biomass thermally deconstructs to a mixture of syngas 

components (CO, H2, CO2, CH4, etc.), tars, and solid char containing residual carbon 

from the biomass and coke deposited on the olivine. Cyclones at the exit of the gasifier 

1 Calcined magnesium silicate, primarily Enstatite (MgSiO3), Forsterite (Mg2SiO3), and Hematite (Fe2O3). A small 

amount of magnesium oxide (MgO) is added to the fresh olivine to prevent the formation of glass-like 

agglomerations formed through biomass potassium interacting with silicate compounds. 

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separate the char and olivine from the syngas. The solids flow to the char combustor 

where the char is burned in air in a fluidized bed, resulting in olivine temperatures greater 

than 1,800°F (982°C). The hot olivine and residual ash is carried out of the combustor by 

the combustion gases and separated using a pair of cyclones. The first cyclone captures 

olivine while the second cyclone captures ash and olivine fines. Hot olivine flows back 

into the gasifier, completing the gasification loop. The hot flue gas from the char 

combustor is utilized for heat recovery and feedstock drying. Ash and olivine fines are 

cooled, moistened to minimize dust, and removed as waste. 

•	 Gas cleanup. Syngas cleanup is defined in this report as reforming of tars, methane, and 

other hydrocarbons followed by cooling, quench, and scrubbing of the syngas for 

downstream operations. The water gas shift reaction also occurs in the reformer. Tars, 

methane, and light hydrocarbons are reformed to syngas in a circulating, fluidized, solid 

catalyst system that resembles a small-scale fluid catalytic cracker (FCC), complete with 

reforming and regeneration operations in separate beds. The syngas is reacted with tar 

reforming catalyst in an entrained flow reactor at a gas hourly space velocity of 

approximately 2,500 h-1. The catalyst is then separated from the effluent syngas in a 

cyclone. From the cyclone, the spent catalyst flows to the catalyst regenerator vessel 

where residual coke from the reforming reactions is removed from the catalyst by 

combustion. The hot catalyst is separated from the combustion flue gas in the regenerator 

cyclone and flows back to the tar reformer reactor to provide the energy necessary for the 

reforming reactions. Additional syngas and unreacted gases from the alcohol synthesis 

reactor are also combusted in the regenerator to provide all the heat necessary for the 

endothermic reforming reactions. A similar approach is documented in a patent 

application by Rentech [24]. The hot reformed syngas is cooled through heat exchange 

with other process streams and scrubbed with water to remove persistent impurities like 

particulates, ammonia, halides, and recalcitrant tars. Scrubber water is purged and treated 

continuously an on-site wastewater treatment facility. After heat recovery, the remaining 

low-quality heat in the flue gas from the catalyst regenerator is utilized for feedstock 

drying. 

•	 Alcohol synthesis. Cooled low-pressure syngas enters a six-stage centrifugal compressor 

system where the pressure is increased to approximately 3,000 psi (207 bar). The 

pressurized syngas is fed to the tube sides of two vertical tubular reactors (shell and tube 

type) operating in parallel. A sulfide-type mixed alcohol catalyst is packed within the 

tubes of the reactor, which is oriented in a down-flow configuration. Reactions at above 

570°F (300°C) convert a portion of the syngas to oxygenate and hydrocarbon products. 

Heat from the reactions is removed by steam generation in the shell side of the tubular 

reactor. The reactor effluent consisting of mixed alcohols, gaseous by-products (such as 

CO2 and methane), and unconverted syngas is cooled through heat exchange with other 

process streams. As the reactor effluent cools, alcohols and water are condensed and sent 

to downstream separation and purification equipment. Unconverted syngas and gas-phase 

by-products flow to an acid gas removal system where a physical solvent (dimethyl ether 

of polyethylene glycol or DEPG) removes most of the H2S and a portion of the CO2. The 

DEPG system requires refrigeration, which is provided by an ammonia-water absorption 

refrigeration system. Absorption refrigeration has the advantage of using heat as the 

major driver instead of electricity used in conventional refrigeration systems. Once the 

12
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

H2S and CO2 are removed, the syngas is recompressed to make up for pressure losses 

through the reactor and acid gas cleanup system, and is recycled to the alcohol synthesis 

reactors. A portion of the unconverted syngas is expanded through a power recovery 

turbine and recycled to the gas cleanup and conditioning section as feed to the tar 

reformer and fuel for the tar reformer catalyst regenerator. The CO2 and H2S from the 

acid gas removal system are further processed in an amine-based acid gas enrichment 

unit and a Merichem LO-CAT sulfur recovery unit where H2S is converted to elemental 

sulfur and stockpiled for disposal. An intermediate CO2-rich stream generated from the 

acid gas enrichment system is recycled to the reformer catalyst regeneration combustor to 

recover heat from residual combustible components in the stream and to prevent the 

emission of hydrocarbons into the atmosphere. 

•	 Alcohol separation. Cooled crude alcohols are de-pressurized and de-gassed in a flash 

separator. The evolved gases are recycled to the gas cleanup section as a feed to the tar 

reformer. The depressurized liquid stream is dehydrated using a molecular sieve system. 

Downstream of the molecular sieve, the dehydrated alcohol stream is introduced to the 

crude alcohol distillation column, which separates methanol and ethanol from higher 

molecular weight alcohols. The overhead stream from the crude alcohol column is 

separated in a second column to a crude methanol stream and an ethanol product that 

meets ASTM fuel ethanol specifications (provided in Appendix J). The methanol 

overhead stream is split and utilized for the following purposes: 

o	 15% (in the base case design) is sent to the molecular sieve system, which is 

adequate to regenerate the molecular sieve beds by flushing the adsorbed water 

[25]. The resulting methanol/water mixture is returned to the gasifier as a source 

of steam for fluidization. 

o	 From 0% to 15% (11% in the base case design) is chilled and mixed with an H2Srich stream from the sulfur recovery system. H2S dissolves in the cold methanol 

and the resulting solution is pumped to the inlet of the alcohol synthesis reactor, 

providing sulfur for catalyst activity management. The amount of H2S is adjusted 

to maintain an average concentration of 70 ppmv at the reactor inlet. It should be 

noted that a range of 0% to 15% is specified because the required amount of H2S 

recycle will depend on the amount of H2S (from sulfur present in the biomass 

feedstock) entering the alcohol synthesis reactor with fresh syngas. 

o	 A small vapor portion (5 wt % in the base case design) containing some methanol 

with gases and volatile compounds is utilized as fuel for the tar reformer catalyst 

regenerator. 

o	 The remaining material (69% in the base case design) is recycled to the alcohol 

synthesis reactor to improve ethanol yields. 

•	 Heat and power. A conventional steam cycle produces steam for direct injection into the 

gasifier and reformer. Indirect steam is used for distillation, absorption refrigeration, and 

the acid gas strippers in the DEPG and acid gas enrichment units. Electricity is generated 

using two steam turbines, with intermediate reheat, to meet the demands of the plant. As 

mentioned earlier in the alcohol synthesis section, additional electricity is generated by 

sending a portion of the pressurized unreacted gases from the alcohol synthesis reactor 

through turbo-expanders. The majority of electricity is used for syngas compression. The 

13
 
 

 

 

 

 

 

 

 

 

steam cycle is integrated with the biomass conversion process. Pre-heaters, steam 

generators, and super-heaters are integrated within the process design to generate the 

steam from boiler feed water. Process condensate is recycled to the steam cycle, degassed, treated, and combined with makeup water. 

•	 Heat integration. Previous analyses of gasification processes have shown the importance 

of properly utilizing the heat from the high temperature streams to improve efficiency of 

the overall thermochemical ethanol plant design [2, 5]. A pinch analysis was performed 

and a heat exchanger network was designed for this ethanol production process. The 

pinch concept offers a systematic approach to optimizing the energy integration of the 

process. Details of the pinch analysis are provided in Section 3.9. 

•	 Cooling water. A cooling water system is included in the Aspen Plus model to determine 

the requirements of each cooling water heat exchanger within the biomass conversion 

process as well as makeup water and power requirements. In a version of the design that 

includes additional water optimization, process condensate is treated and combined with 

the cooling water makeup to reduce fresh water consumption. 

2.3 Aspen Plus Model 

An Aspen Plus Version 2004.1 simulation was used as the basis for this report. The gaseous and 

liquid components that are important for the process were included as distinct molecular species 

using Aspen Technology’s component properties database. The biomass feedstock, ash, and char 

components were modeled as non-conventional components. The thermochemical ethanol design 

requires processing of three different phases of matter (solid, liquid, and gas phases). Therefore, 

no single physical property package was sufficient to describe the entire plant from feed to 

products. Different property packages were used within the Aspen Plus simulation to more 

accurately represent chemical component behaviors in specific process areas. The Redlich– 

Kwong–Soave with Boston–Mathias correction (RKS-BM) equation of state was used 

throughout much of the process simulation for high temperature, high pressure gas and liquid 

phase behavior. The non-random, two-liquid (NRTL) option with ideal gas properties was used 

for alcohol separation calculations. The ASME 1967 steam table correlations were used for the 

steam cycle calculations. The ELECNRTL package was used to model the electrolyte species in 

the quench water system. The physical solvent absorption system and the ammonia-water 

absorption refrigeration system were modeled in separate simulations and their results were 

incorporated into the main model. The perturbed-chain statistical associating fluid theory (PCSAFT) package was used to model physical solvent (DEPG) absorption processes in a separate 

simulation. 

14
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3 Process Design and Cost Estimation 

The process design for the thermochemical ethanol production facility is based on current data 

and R&D performance goals for catalytic reforming and alcohol synthesis. The design therefore 

presents the anticipated results of achieving those technical goals. The specific technical goals 

are presented in Appendix I. 

The process design broadly consists of the following areas: 

• Feed handling and drying 

• Gasification 

• Gas cleanup 

• Syngas compression, alcohol synthesis, and acid gas removal 

• Alcohol separation 

• Integrated steam system and power generation cycle 

• Cooling water and other utilities 

The following sections present process overviews, design bases, and installed equipment cost 

estimations for each process area of the plant. Heat exchanger equipment sizes and costs are 

discussed separately in Section 3.9. 

3.1 Area 100: Feed Handling and Drying 

The following section presents an overview of the feed handling and drying facilities. The basis 

for the system design is discussed, and cost estimates for the feedstock are provided based on 

capital requirements and operating costs of the facilities. The information provided is based on 

work done at the Idaho National Laboratory (INL). 

3.1.1 Area 100 Overview 

This area of the process accommodates the delivery of biomass feedstock from the field to the 

biorefinery, short term on-site storage, and the preprocessing/preparation of the feedstock for 

conversion in the gasifier. The design is based upon a specific woody biomass feedstock – 

southern pine trees. Process flow diagrams (PFD) for the feed handling and drying area are 

shown in PFD-P820-A101 and PFD-P820-A102 in Appendix E. A simplified flow diagram for 

the woody biomass feedstock supply system is provided in Figure 5. 

Standing southern pine trees are cut with an average diameter at breast height (DBH) of 8–10 

inches using a feller buncher with an accumulator arm. Material is harvested 8 hours per day, 5 

days per week, and 50 weeks per year. The cut trees are piled on the ground and transpirationally 

dried from 50% moisture content (MC) to approximately 35% MC, wet basis [26–29]. A grapple 

skidder drags the piled trees to the landing, and the trees are piled near the flail chain debarker. 

The flail chain debarker is equipped with a grapple on the end that loads material from the deck 

into the flail chain debarker. Approximately 20% of the material is ejected during this step. The 

ejected material is bark, tree tops, and limbs. The debarked, delimbed logs are fed into the 

chipper, which uses a 2-inch internal screen. Pneumatics may be integrated into the chipper, 

which would increase grinder efficiency and provide additional drying beyond that which is 

15
 
 

 

 

 

 

accomplished in the absence of pneumatics. Chips are ejected from the chipper directly into a flat 

floor chip trailer, and once the truck reaches capacity the material is taken to the biorefinery. 

Figure 5. Woody biomass feedstock supply system 

Note: Only biomass material streams are shown for simplification. See Appendix F for full schematic. 

To increase material availability to the system, reduce dry matter losses, and meet cost targets, 

some of the limbs are fed through the chipper in a stream separate from the clean chips. This 

slash stream, as we denote it, is passed through a trommel screen to sift out high-ash dirt and 

bark. However, it is assumed that some of the ash and bark have already been separated from the 

limbs and tops during the high-impact flail process. The cleaned slash stream is also ejected into 

the back of a truck and transported to the biorefinery. At the biorefinery, the trucks are weighed 

and a truck dumper unloads the truck contents into a hopper. A dust collection system is present 

during unloading to prevent excessive dust accumulation and to limit dry matter losses. The 

chips are cleaned using an electromagnet, moisture is monitored, and material is passed over a 

vibratory conveyor to remove any excess dirt/ash that may still be present. The material is then 

conveyed into a pile using a circular stacker and an overpile reclaimer. A five-day supply of 

material is stored at the refinery. The pile, which acts as a queue, is on an asphalt pad with a long 

grate running through it that allows material to flow out of the pile, and up to an additional 5% 

moisture loss results from forced-air ambient air drying. A front end loader continuously pushes 

material onto the grate to maintain flow. A conveyor located at the bottom of the pit under the 

grate conveys the biomass into a waste heat dryer, where the biomass is dried to 10% moisture 

content using hot flue gas from the biomass conversion processes. From the dryer, the biomass is 

conveyed into a metering bin where it is fed into the conversion process. 

16
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.1.2 Area 100 Design Basis 

Table 5 presents the design basis for the feed handling and drying area of the plant. The primary 

objective of this equipment is to dry the biomass feed from 35 wt % moisture content to 10 wt % 

moisture content at the inlet to the gasifier in order to maximize the efficiency of the gasifier and 

minimize the required size of downstream equipment. 

Table 5. Design Basis for Feed Handling and Drying Area 

Biomass feed type Wood from southern pine trees 

Biomass feed rate (dry basis) 2,000 metric tonnes per day 

Biomass feed moisture content 35 wt % 

Additional ambient drying at plant 5 wt % 

Target biomass moisture content to gasifier 10 wt % 

Assuming biomass delivery by truck with a per truck capacity of 25 tons [30] of biomass at 35 

wt % moisture content, approximately 136 truck deliveries per day would be required to operate 

the plant at design capacity of 2,000 metric tonnes per day (2,205 tons per day). 

3.1.3 Area 100 Equipment Cost Estimations 

INL, in partnership with other national labs, has developed a woody supply system design and 

modeled the costs using a woody biomass supply system model. The model is continuously 

refined through input from industry, academia, and national labs to improve the accuracy of the 

model parameters and output. Future model iterations will identify key sensitivities in the system 

in terms of costs, material flows, and equipment performance parameters. 

Table 6 shows the costs associated with the operations in the supply system (Figure 5). The data 

presented include unit operation costs for the design case scenario for the supply of woody 

biomass (pulpwood-size southern pine) to a gasification process. All values are presented in units 

of 2007 U.S. dollars per dry ton of feedstock. 

17
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 6. Unit Cost Breakdown for Woody Biomass Feedstock (Southern Pine) 

Unit Operation Cost Contribution 

($/dry U.S. ton) 

Total feedstock cost through gasifier reactor inlet $61.57 

Total feedstock logistics $46.37 (harvest through insertion to gasifier reactor inlet)
 

Total cost of grower payment $15.20
 

Harvest and collection $18.75 

Capital costs $5.60 

Operating costs $13.15 

In-field preprocessing $11.42 

Capital costs $4.20 

Operating costs $7.22 

Transportation and handling $8.95 

Capital costs $2.95 

Operating costs $6.00 

In-plant receiving, storage, queuing, feed preprocessing $7.25 

Capital costs $2.10 

Receiving and handling costs $2.10 

Other operating costs $3.05 

The total delivered cost of woody biomass from pulpwood-size pine trees, as determined from 

the woody biomass supply system model, is $46.37/dry U.S. ton (short ton). When the grower 

payment of $15.20/dry U.S. ton is added, the total feedstock cost through the gasifier inlet is 

$61.57/dry U.S. ton. This figure includes ownership costs and operating costs. 

The supply system model incorporates a combination of values and relationships obtained from 

other national laboratories, publications, consultation with academia and staff from the U.S. 

Department of Agriculture and the U.S. Forest Service, and published and unpublished INL data. 

Further details on the model are provided in Appendix F. 

Equipment lists for the feed handling and drying area, as well as other pertinent information, are 

provided in Appendix A and Appendix B. All purchased and installed equipment costs for this 

area are shown as zero for the plant economics because all of these capital costs are included in 

the delivered feedstock cost. 

3.2 Area 200: Gasification 

The following section presents an overview, basis for design, and cost estimates for construction 

of the gasification facilities. 

18
 
19 

3.2.1 Area 200 Overview 

This area of the process converts a mixture of dried feedstock (10 wt % moisture) and steam to 

syngas and char. Heat is provided indirectly through circulation of olivine (heat transfer media) 

that is heated by the combustion of char downstream of the gasifier. Steam is injected into the 

gasifier to serve as a fluidizing medium and reactant at the high gasifier temperatures. A 

simplified flow diagram of the gasification process area is provided in Figure 6. 

Figure 6. Simplified process flow diagram of gasification area 

After going through the feed handling and drying area, the dried wood chips enter the 

gasification area, as shown in PFD-P820-A201 in Appendix E. The design includes two parallel 

gasifier trains, each with a biomass processing capacity of 1,000 dry metric tonnes per day. The 

gasifier (R-201) used in this analysis is a low-pressure indirectly-heated circulating fluidized bed 

(CFB) gasifier. The gasifier output was modeled using correlations based on run data from the 

Battelle Columbus Laboratory (BCL) 9 tonne/day test facility. These correlations are provided 

for reference in Appendix G. 

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

vessel and the char combustor (R-202). In this case the medium is synthetic olivine – calcined 

magnesium silicate consisting of Enstatite (MgSiO3), Forsterite (Mg2SiO3), and Hematite 

(Fe2O3). A small amount of magnesium oxide (MgO) is added to the fresh olivine. The MgO 

rejects the potassium present in biomass as ash by forming a high melting (~2,370°F/1,300°C) 

ternary eutectic with the silica, thus sequestering it. Without MgO addition, the potassium will 

form glass (K2SiO4) by interacting with the silica in the system. K2SiO4 has a low melting point 

Biomass

Feed

Gasifier

Combustor

Cyclones

Primary 

Cyclone

Olivine Olivine + 

Char

Synthesis 

Gas

Flue

Gas

Steam Air

Secondary

Cyclone

Ash & 

Olivine 

Fines
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

(~930°F/500°C) and its formation would cause the bed media to become sticky, agglomerate, 

and eventually defluidize. Potassium carry-over from the gasifier to downstream processes is 

also significantly reduced with the addition of MgO. The biomass ash is assumed to contain 

0.2 wt % potassium. The MgO molar flow rate is set at 2 times the molar flow rate of potassium. 

With an expected potassium inlet flow rate of 0.087 lb-moles per hour, the estimated MgO 

requirement is 0.174 lb-moles per hour or 7.0 lb per hour. 

The fluidization medium in the gasifier is steam supplied from the steam cycle (area 600). The 

steam-to-feed ratio is set to 0.4 lb of steam per lb of bone dry biomass feedstock. The gasifier is 

designed to operate at 18 psig to allow for pressure drops in subsequent equipment so that no 

further compression is required until after gas cleanup and quench. The average olivine 

circulation rate is approximately 27 lb of olivine per lb of bone dry wood. Fresh olivine is added 

to compensate for losses from the gasifier and combustor cyclones. 

The operation of the gasifier and char combustor resembles that of a fluid catalytic cracker 

(FCC), a common oil refinery unit operation. Like an FCC, the gasifier and char combustor will 

remain in heat balance by producing (gasifier) and burning (char combustor) enough char to 

satisfy the energy requirements of the operation. The amount of char formed in the gasifier is an 

inverse function of temperature. Unless supplemental fuel is provided to the char combustor for 

temperature control, the gasifier and combustor reach thermal equilibrium based on the amount 

of char formed in the gasifier. If the gasifier temperature is lower than the equilibrium 

temperature, then more char is formed and more heat is generated by char combustion. This 

results in more heat transfer from the combustor to the gasifier, thus bringing the gasifier 

temperature up toward equilibrium. Although the gasifier, char combustor, and associated piping 

will be insulated to maximize thermal efficiency, some heat losses will occur. Two percent (2%) 

of the lower heating value of the feed biomass (dry basis) is assumed to be lost from the system. 

The resulting gasifier temperature is 1,596°F (869°C) and the char combustor temperature is 

1,808°F (987°C). The char combustor is operated with 20% excess air (based on the 

stoichiometric requirement) to ensure complete oxidation of char components. 

Olivine is recovered from the raw syngas exiting the gasifier using two stages of cyclone 

separators. Nearly all of the olivine and char (99.9% of both) are separated in the primary 

gasifier cyclone (S-201) and fed by gravity to the char combustor. A secondary cyclone (S-202) 

removes 90% of residual fines. 

The combustion flue gas and olivine from the char combustor are sent to a cyclone separation 

system similar to the gasifier cyclone. A primary combustor cyclone (S-203) separates the 

olivine (99.9%) from the combustion gases and the olivine is gravity-fed back to the gasifier. 

The flue gas from the primary cyclone is sent through a secondary cyclone (S-204) and then 

utilized for heat recovery before it is mixed with flue gas from the tar reformer area. The 

combined flue gas stream is utilized for feedstock drying and finally enters the flue gas scrubber 

for removal of residual particulate matter and regulated chemical species such as SO2. 

The mixture of olivine, ash, and catalyst fines from the secondary flue gas cyclone is indirectly 

cooled to 300°F in the water-cooled screw conveyor (M-201). Water is added directly to the 

solids for further cooling and dust prevention. The solids are disposed as landfill. 

20
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Another similarity to an FCC is the importance of pressure balance between the gasifier and char 

combustor in reliably controlling the circulation of olivine through the system. The pressure 

balance ensures that olivine circulation is maintained from the gasifier to the char combustor and 

back to the gasifier (not vice versa). The pressure balance in such systems is usually controlled 

by two slide valves (or L-valve loop seal or J-leg loop seal), one controlling the flow of spent 

olivine from the gasifier cyclones to the char combustor and the other controlling the flow of 

regenerated olivine from the primary char combustor cyclone to the gasifier. The pressure 

balance for these types of processes is addressed during the detailed design phase of a project. 

3.2.2 Area 200 Design Basis 

The design basis for the gasification area is presented in Table 7. This basis reflects the overall 

plant capacity of 2,000 dry metric tonnes per day (2,205 dry tons per day). 

Table 7. Design Basis for Gasification Area 

Parallel gasification trains Two (2) 

Biomass feed rate per train (dry basis) 1,000 metric tonnes per day 

Biomass moisture content to gasifier 10 wt % 

Gasifier Char combustor 

Operating pressure (psia) 33.0 29.0 

Operating temperature (°F) 1,596 1,808 

3.2.3 Area 200 Equipment Cost Estimations 

Capital costs for the equipment in this area were estimated by Taylor Biomass Energy for a 500 

metric tonnes per day (MTPD) unit. The estimate included the cost for a thermal tar reformer, 

which will be discussed in the following section. While the Taylor Biomass Energy base design 

is for a 500 MTPD modular unit, the company supports the feasibility of scaling the unit to a 

1,000 MTPD capacity using a scaling exponent of 0.6 as shown below. 

1,000 MTPD equipment cost = 500 MTPD equipment cost x (1,000 / 500) 0.6 

1,000 MTPD equipment cost = $9.7 MM x 1.5157 = $14.7 MM 

Because the total capacity of the thermochemical plant design is 2,000 MTPD, two parallel 1,000 

MTPD units will be incorporated in the design. The cost estimate from Taylor Biomass Energy is 

summarized in Table 8 along with scaled costs for two parallel 1,000 MTPD units. Taylor 

Biomass Energy’s quotation for the 500 MTPD unit is more representative of a first-of-its-kind 

plant than an nth-plant, suggesting that the capital cost estimate is conservative. However, the 

design presented in this report has minor variations from the Taylor Biomass Energy design [31]. 

It is assumed that additional costs for these variations are adequately covered by the 

conservatism in the Taylor Biomass Energy quotation, as nth-plant costs are expected to be 10% 

to 15% lower than costs for first-of-a-kind predecessors. Because the quote also includes gas 

cleanup equipment (thermal cracker or tar reformer), two-thirds of the total equipment costs is 

allocated to the two 1,000 MTPD gasifiers and the remaining one-third is allocated to the tar 

reformer (cost details provided in Appendix B). The installation factor applied to calculate total 

installed cost (TIC) was determined from the Taylor Biomass Energy quotation. 

21
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 8. Estimates for Gasification and Tar Reforming Facilities 

Taylor Biomass 

Energy 500 MTPD 

Single Train 

Scaled 

1,000 MTPD 

Single Train 

Two Scaled Parallel 

1,000 MTPD Trains 

(Aug 2010) (Aug 2010) (Aug 2010) (2007) 

Dry feedstock capacity 

(metric tonnes per day) 500 1,000 2,000 2,000 

Total purchased equipment cost 

(TPEC) $9.7 MM $14.7 MM $29.4 MM $28.5 MM 

Equipment installation factor 2.31 2.31 2.31 2.31 

Total installed cost (TIC) $22.4 MM $33.9 MM $67.9 MM $64.9 MM 

TIC allocated to gasifier (2/3) $43.3 MM 

TIC allocated to tar reformer (1/3) $21.6 MM 

The Taylor Biomass Energy system includes the following equipment: 

• Biomass feeding system (excluding storage, pre-treatment, and transport equipment) 

• Gasification process reactors (gasification, gas conditioning, and combustion) 

• All cyclones and interconnecting ducting 

• Combustion air blower 

• Combustion air pre-heat exchanger(s) 

• Start-up equipment 

• Syngas polishing scrubber (final cleanup and particulate removal) 

• Ash discharge system 

• Emergency flare system 

• Process instrumentation 

• Olivine and magnesium oxide makeup system 

• Inert gas system. 

Table 9 presents the anticipated equipment requirements and costs for the gasification area. The 

gasifier cost presented in this table reflects the two-thirds cost assumption explained above. 

Table 9. Cost Estimate for Gasification Area 

Equipment Description TPEC TIC 

(2007 $k) (2007 $k) 

Gasifier (ancillaries included) 18,744 43,254 

Area A200 subtotal 18,744 43,254 

22
 
 

 

 

 

 

 

 

 

 

 

3.3 Area 300: Gas Cleanup 

The following section presents an overview, basis for design, and cost estimates for construction 

of the gas cleanup area. 

3.3.1 Area 300 Overview 

This section of the process cleans and cools the syngas so it can be compressed and converted 

into alcohols. Figure 7 presents a simplified process flow diagram for the gas cleanup area of the 

plant, which includes the following major steps: 

• Tars, methane, and other hydrocarbons in the syngas are reformed to CO and H2. 

• Particulates and other contaminants are removed by a quenching and scrubbing process. 

Synthesis 

Gas from 

Gasifier 

Tar Reformer 

Cyclone 

Tar 

Reforming 

Catalyst 

Circulation 

Synthesis 

Gas 

Air 

Flue 

Gas to Heat 

Recovery 

Unconverted 

Synthesis 

Gas Recycle 

Catalyst

Regenerator 

Cyclone 

Steam 

Purge 

Water to 

Treatment 

Water 

Circulation Venturi Scrubber 

Syngas Conditioning by Tar Reforming 

Make-Up 

Water 

Syngas Quench & Scrubbing 

Scrubbed 

Synthesis Gas Cyclone Separator 

Figure 7. Simplified process flow diagram of gas cleanup area 

The type and the extent of cleanup are dictated by the requirements of the alcohol synthesis 

catalyst and other downstream equipment. This design differs from the previous design [2] in 

that sulfur removal is not performed on the syngas upstream of the alcohol synthesis reactor 

(detailed discussion in Section 3.4). 

The syngas from the secondary gasifier cyclone is sent to the catalytic tar reformer (R-303 in 

PFD-P820-A303, Appendix E), which is an entrained-flow, fluidized catalytic reactor. In the 

reformer, methane, tars, and other hydrocarbons are converted to CO and H2, and NH3 is 

converted to N2 and H2. Because the reforming catalyst has significant water-gas-shift activity, 

the syngas is assumed to exit the reformer at water-gas-shift equilibrium. In the Aspen Plus 

simulation, the conversion of each compound is set to match targets that are believed to be 

23
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

attainable through near-term research efforts, as summarized in Table 10. Similar to the 

gasification area, the heat for tar conversion is supplied by circulating catalyst (Ni/Mg/K 

supported on alumina) [32] between the tar reformer reactor and the catalyst regenerator. The 

catalyst regenerator burns carbon (coke) deposits off the catalyst particles, regenerating the 

catalyst activity and providing heat for the tar reforming reactions. The concentration of coke 

precursors in the syngas feed to the tar reformer will vary depending on feedstock type and 

conversion in the gasifier. Because the rate of coke formation in the tar reformer is not sufficient 

to sustain the endothermic reforming reactions, supplemental combustion gases are supplied to 

the catalyst regenerator from various process areas. Air is supplied to the catalyst regenerator by 

the combustion air blower (K-305) to provide the oxygen for combustion. Continuous operation 

has been demonstrated with simulated syngas (including some H2S) using NREL’s reforming 

catalyst [24]. The raw syngas from the gasifier and supplemental steam serve as fluidization 

media for the catalyst in the reformer. No significant amount of supplemental steam is required 

in the base case design, as the steam present in the syngas from the gasifier is nearly sufficient to 

maintain a target molar steam-to-carbon ratio of 2. The spent catalyst from the tar reformer is 

separated from the syngas in a two-stage cyclone separator (S-306) and flows to the catalyst 

regenerator (R-301A). The base tar reformer operating temperature for this design is 1,670°F 

(910°C) at the reactor outlet. This temperature is based on experimental results using NREL’s 

reforming catalyst, which showed that the catalyst maintains high reforming activity in the 

presence of biomass-derived syngas at 1,697°F (925°C). The lower temperature threshold of the 

high performance has not yet been determined. The model can be modified to maintain the 

reformer at 1,697°F instead of 1,670°F without significant changes in yields. 

The combustion flue gas and catalyst from the regenerator are sent to a cyclone separator system. 

A two-stage cyclone (S-307) separates the regenerated catalyst from the combustion gases and 

the catalyst flows back to the tar reformer. The flue gas is then utilized for heat recovery before it 

is mixed with flue gas from the gasification area. The combined flue gas stream is used for the 

feed drying system and then sent to a flue gas scrubber. 

Hot syngas from the tar reformer is cooled to 140°F (60°C) through heat exchangers, which are 

used to heat other process streams. A syngas scrubbing system removes impurities such as 

particulates, ammonia, halides, and residual tars. The scrubbing system consists of a venturi 

scrubber (M-302), cyclone separator (M-301), and a quench water circulation system. The 

quench water is contained in a closed recirculation loop with heat rejected to the cooling tower 

through the quench water cooler (H-301). The quench water circulation rate is set to an L/G 

(liquid flow to actual volumetric gas flow) ratio equal to 1 liter/m3

. This ratio is based on a study 

of venturi scrubbers that suggests operating conditions for near-optimal performance [33]. The 

syngas exits the quench system at a temperature of 130°F (54°C). A continuous purge and fresh 

water makeup is maintained in the circulation loop at approximately 3 gpm to control the solids 

level in the circulating water. The purge stream is sent to a clarifier to separate the solids from 

the water. The solids are routed for disposal and the water, which includes process condensate, is 

recycled to the reverse osmosis (RO) purification system in the base case design. In a version of 

the design that includes water optimization, the water is treated and used for flue gas scrubbing 

and the remainder is mixed with the cooling water makeup stream (shown in PFD-P820-A701 in 

Appendix E). 

24
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.3.2 Area 300 Design Basis 

Table 10 summarizes the 2012 targets for hydrocarbon conversions during continuous operation. 

While most of these targets have been demonstrated at NREL at bench scale, future efforts are 

focused on continuous operation at the pilot scale for extended periods. Equipment is currently 

being installed at NREL for pilot-scale demonstrations. 

Table 10. Target Performance of Reforming Catalyst 

Compound 2012 Target Conversion 

Methane (CH4) 80 % 

Ethane (C2H6) 99 % 

Ethylene (C2H4) 90 % 

Tars (C10+) 99 % 

Benzene (C6H6) 99 % 

Ammonia (NH3) 90 % 

These target conversions were incorporated into the Aspen Plus model to simulate expected tar 

reformer performance. The compositions of the reformer inlet and outlet streams are presented in 

Table 11. 

Table 11. Reformer Conditions, Gas Compositions, and Reactor Basis for the Base Case 

Inlet Outlet 

Temperature, °F (°C) 1,632 (889) 1,670 (910) 

Pressure, psia 28.7 25.7 

H2:CO molar ratio 0.864 1.247 

Gas composition (mole %) Wet Dry Wet Dry 

Hydrogen (H2) 19.7 27.6 40.3 48.1 

Carbon dioxide (CO2) 15.5 21.7 9.4 11.2 

Carbon monoxide (CO) 22.8 32.0 32.3 38.5 

Water (H2O) 28.5 --­ 16.2 --­

Methane (CH4) 9.3 13.0 1.5 1.8 

Acetylene (C2H2) 0.1 0.2 115 ppmv 138 ppmv 

Ethylene (C2H4) 1.6 2.3 1,295 ppmv 1,546 ppmv 

Ethane (C2H6) 1,608 ppmv 2,250 ppmv 13 ppmv 15 ppmv 

Benzene (C6H6) 449 ppmv 629 ppmv 3.6 ppmv 4.3 ppmv 

Tars (C10H8) 822 ppmv 1,150 ppmv 6.6 ppmv 7.9 ppmv 

Ammonia (NH3) 1,158 ppmv 1,621 ppmv 93 ppmv 111 ppmv 

Hydrogen sulfide (H2S) 133 ppmv 186 ppmv 106 ppmv 127 ppmv 

Nitrogen (N2) 1,539 ppmv 2,154 ppmv 1,647 ppmv 1,966 ppmv 

Reactor space velocity (h-1) 2,476 

Reactor volume (ft3

) 2,713 

25
 
 

 

 

 

 

 

 

 

 

 

 

 

 

3.3.3 Area 300 Equipment Cost Estimations 

Table 12 presents the equipment list and cost estimates for the gas cleanup area of the plant. The 

tar reformer cost presented in this table reflects the one-third cost assumption based on Taylor 

Biomass Energy’s cost estimate for gasification and gas cleanup, as discussed in detail in Section 

3.2.3. Equipment costs in the table were obtained from quotations by Taylor Biomass Energy (tar 

reformer) and TurboSonic (syngas quench and scrubbing system). 

Table 12. Cost Estimate for Gas Cleanup Area 

Equipment Description TPEC TIC 

(2007 $k) (2007 $k) 

Tar reformer and syngas scrubber 11,697 24,330 

Heat integration 929 2,610 

Area A300 subtotal 12,626 26,940 

3.4 Area 400: Alcohol Synthesis, Syngas Compression and Acid Gas Processing 

The following section presents an overview, basis for design, and cost estimates for construction 

of the alcohol synthesis reactor, syngas and recycle compressors, turbo-expander, and acid gas 

removal facilities. 

3.4.1 Area 400 Overview 

The major function of this area is synthesis of mixed alcohols from syngas by catalytic 

conversion. The following major steps occur in the alcohol synthesis area: 

•	 Syngas compression to 3,000 psia 

•	 Alcohol synthesis 

•	 Steam generation from cooling of the exothermic alcohol synthesis reaction 

•	 Cooling and recovery of alcohols from reactor effluent (in the liquid phase) 

•	 CO2 and H2S removal 

•	 Partial recycle of H2S into the reactor by dissolving in chilled methanol 

•	 Sulfur recovery from removed H2S 

•	 Recycle of most of the cleaned, unreacted syngas to the reactor after recompression 

•	 Power recovery from gaseous effluents in a turbo-expander. These gases are then sent to 

the reformer and reforming catalyst regenerator. 

Figure 8 presents a simplified process flow diagram for the alcohol synthesis area. 

26
 
 

 

 

 

Steam Drum 

From Synthesis 

Gas Compression 

Boiler Feed 

Water 

Steam 

Sulfur 

To Alcohol 

Separation 

Recycle 

Compressor 

Low 

Pressure 

Recycle 

Synthesis Gas 

Recycle to Tar 

Reformer 

TurboExpanders 

Syngas Compression for Alcohol Synthesis 

1st Stage 

22.7 PSIA 

4th Stage 

Alcohol Synthesis, Syngas Recycle, Acid Gas Removal & Sulfur Recovery 

Scrubbed 

Syngas from Gas 

Cleanup and 

Conditioning 

Flash 

Drum 

2nd Stage 3rd Stage 

5th 6 Stage th Stage 

3,000 PSIA 

Compressed 

Synthesis Gas 

CO2 

Acid Gas 

Removal & 

Sulfur Recovery 

Figure 8. Simplified process flow diagram of alcohol synthesis area 

The scrubbed syngas is compressed to 3,000 psi (207 bar) using a six-stage centrifugal 

compressor with inter-stage cooling as shown in Figure 8 and PFD-P820-A401 (Appendix E). 

The compressor was modeled such that each section has a polytropic efficiency of 78% to 87% 

(based on vendor quotes), assumed mechanical efficiency of 98%, and intercooler outlet 

temperatures of 110°F (43°C). Each inter-stage cooler includes water, air, and process stream 

27 
 

 

 

 

 

 

 

 

 

 

 

 

exchangers. These exchangers are deliberately oversized to minimize pressure drops to an 

average of 6 psi (0.41 bar) per inter-stage (i.e., an average of 2 psi pressure drop in each of the 

three inter-stage coolers). 

Minimizing pressure drop in the coolers will minimize electrical consumption as the compressors 

represent the most significant electrical load on the system, consuming approximately 47 MW of 

electricity in the base case. 

The compressed fresh syngas is mixed with recycled syngas and methanol, and pre-heated to 

595°F (313°C) before entering the alcohol synthesis reactor. Within the reactor, the syngas 

contacts a metal-sulfide catalyst. For this report, performance was modeled after The Dow 

Chemical Company’s potassium-promoted cobalt molybdenum sulfide catalyst. This catalyst is 

described in detail in U.S. Patent 4,882,360. 

The overall stoichiometric reaction for alcohol synthesis is summarized in Equation 4: 

n CO + 2n H2  CnH2n+1OH + (n-1) H2O (Eq. 4) 

Light hydrocarbons, methyl esters, and aldehydes are produced in smaller quantities through 

similar chemical routes. The stoichiometry suggests an optimum H2:CO ratio of 2.0. However, 

sulfide-type mixed alcohol catalysts have significant water-gas-shift activity and will generate H2 

and CO2 from CO and H2O as shown in Equation 5. For this reason, an H2:CO ratio of less than 

2.0 is preferred for this design. 

CO + H2O  H2 + CO2 (Eq. 5) 

A kinetic model, supplied to NREL by The Dow Chemical Company under confidentiality 

agreement, was used to model alcohol production over the promoted sulfide catalyst. The model 

was developed at The Dow Chemical Company using both bench-scale and pilot-scale data, and 

it shows agreement with recent experimental results, as shown in the parity plots in Figure 9. It 

should be noted that the model is dynamic and incorporates effects of reactant and product 

concentrations, reactor pressure, catalyst particle size, reactor geometry, heat transfer 

coefficients, and temperature. The predictions in the full process model are robust and represent 

a significant improvement in this design report relative to earlier predictions. 

28
 
 

 

 

 

 

 

 

 

 

 

 

 

METHANOL Productivity ETHANOL Productivity PROPANOL Productivity METHANE Productivity 500 

Data 

x = y	 

500 

Data 

x = y	 

250 

Data 

x = y	 

200 

400 400 200 

150 

300 300 150 

100 

200 200 100 

50 

100 100 50 

Data 

x = y 

Model 

0 0 0 0 

0	 100 200 300 400 500 0 100 200 300 400 500 0 50 100 150 200 250 0 50 100 150 200 

Experiment 

Figure 9. The Dow Chemical Company model predictions compared to experimental data 

Notes: 1) Unit for productivity: g/kg catalyst/h. 2) Experimental data spans a wide range of operating temperatures, 

pressures, and feed compositions.
 

Other mixed alcohol catalysts may fit into this process configuration. However, the breadth of 

information available for sulfide-type catalysts, coupled with their relative cost and productivity 

and the availability of The Dow Chemical Company’s kinetic model, warranted significant 

consideration of metal sulfide-type catalysts. Further, because The Dow Chemical Company and 

NREL signed a cooperative research and development agreement (CRADA) in 2009, any 

demonstrated or anticipated improvements to catalyst performance can be incorporated into The 

Dow Chemical Company’s kinetic model and reflected in future State of Technology (SOT) 

reports. 

Heat must be removed from the reactors because the synthesis reaction is exothermic. 

Temperature is controlled by steam generation on the shell side of the reactor. Steam is generated 

at 1,336 psia and 581°F (305°C) and maintained at these conditions by back-pressure control at 

the outlet of the steam drum (T-605). 

The effluent gases from the reactor are cooled and flashed to remove a liquid stream rich in 

alcohols and other minor liquid components, which are sent for purification in area 500. The 

gaseous stream is cleaned to remove CO2 and H2S, as explained below, and recycled to the 

reactor. In the base case design, 12% of the recycle gas is purged to prevent excessive buildup of 

inert gases in the synthesis loop and expanded through a turbo-expander to recover power. After 

power recovery, purge gases are split between the reformer (75%) for carbon recovery and the 

reformer catalyst regenerator (25%) for heat generation. The fraction of the recycle split impacts 

the gas composition at the reactor inlet, especially with respect to methane and other 

hydrocarbon content. 

A major difference in this design compared to previous designs is the location of the acid gas 

removal system. In the previous design [2], the amine-based acid gas removal system was 

located upstream of the alcohol synthesis reactor. In this design, the acid gas removal system 

resides within the alcohol synthesis reactor recycle loop. The new location was chosen due to 

several considerations: 

•	 CO2 control – The CO2 concentration in the alcohol synthesis reactor can be maintained 

at low levels without necessitating a large purge from the recycle stream, which improves 

29
 
 

 

 

 

 

 

reactor performance and overall carbon efficiency. The compression energy requirement 

for recycled syngas is low because the stream is already at high pressure within the 

recycle loop. This was the primary consideration for choosing the current acid gas 

removal design and location. 

•	 Metal-sulfide alcohol synthesis catalyst – To maintain activity and selectivity, sulfide 

catalysts require a small co-feed of reactive sulfur and it is, therefore, not necessary to 

remove sulfur from the syngas stream prior to the reactor. 

•	 Variability of feedstock sulfur – By placing the acid gas removal and sulfur reinjection 

operations within the alcohol synthesis recycle loop, the CO2 and H2S concentrations in 

the alcohol synthesis reactor and recycle can be more effectively controlled. This allows 

the flexibility of processing feedstocks with varying sulfur content without process 

modifications. 

•	 Sulfur containment – With sulfur removal operations within the syngas recycle loop, 

purge gases are also cleaned of sulfur, thus reducing gaseous sulfur emissions from the 

plant and eliminating recycle of sulfur to the reformer. 

H2S/CO2 removal and sulfur management systems represent a substantial portion of the alcohol 

synthesis area of the plant. Figure 10 presents a simplified schematic of these systems, and 

details of the systems are provided in PFD-P820-A403 through PFD-P820-A406 in Appendix E. 

30
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Acid Gas to 

Sulfur Recovery 

MeOH from 

Alcohol 

Separation 

H2S Recycle to MAS by Methanol 

H2S & MeOH 

Recycle to 

MAS Reactor 

Absorber 

Absorber 

Unreacted 

Gases from 

MAS Reactor 

Recycle to 

MAS Reactor 

Recycle to Tar 

Reformer / Fuel 

Combustor 

H2S + CO2 

to Acid Gas 

Enrichment 

Flash 

H 2 S Stripper 

H2S & CO2 Removal by DEPG Physical Solvent Absorber 

To Tar Reformer to 

Catalyst Regenerator 

Acid Gas 

to MeOH 

Recycle 

& Sulfur 

Recovery 

Regenerator 

Acid Gas Enrichment by Selective Amine 

Sulfur Recovery by Merichem LO-CAT 

CO2 Vent 

Venturi

Oxidizer 

Sulfur 

Slurry 

Air 

Absorber 

Figure 10. Simplified flow diagram of the H2S/CO2 removal and sulfur recovery 

The H2S/CO2 removal and sulfur management systems include the following: 

H2S and CO2 removal (M-401) – This system utilizes a physical solvent, dimethyl ethers of 

polyethylene glycol (DEPG), to remove nearly all of the H2S and a portion of the CO2 from the 

alcohol synthesis reactor effluent. The circulation rate of the DEPG solvent can be adjusted to 

control the removal rate of CO2 from the stream. Because a significant portion of the effluent 

stream is recycled to the alcohol synthesis reactor, control of CO2 removal allows control of CO2 

in the synthesis reactor inlet. The outlet from the DEPG absorber is split into two streams. A 

portion is recycled to the synthesis reactor (R-410) inlet and the remainder is sent through two 

expanders (K-412A and K-412B) and recycled to the tar reformer/fuel combustor. The DEPG 

from the absorber bottom is sent to a lower pressure flash drum to selectively expel less soluble 

compounds such as hydrogen, carbon monoxide, and methane. The 250 psia flash stream is sent 

to the low pressure expander (K-412B) and recycled to the tar reformer/fuel combustor. The 

DEPG solvent is then stripped of the remaining CO2 and H2S, chilled, and circulated back to the 

absorber. The stripper off-gas stream, rich in CO2 and H2S, is sent to acid gas enrichment (M403). 

Ammonia absorption refrigeration system (M-402) – This system provides 30°F chilling 

capacity for the DEPG acid gas removal system (M-401) and the methanol/H2S absorption 

system (M-404). The unit utilizes ammonia absorption technology offered by Colibri–BV. 

31
 
 

 

 

 

 

 

 

 

 

 

 

Acid gas enrichment system (M-403) – This system utilizes a selective amine to increase the 

H2S concentration of the acid gas from M-401 for downstream processes. The selective amine 

absorbs nearly all of the H2S entering from the DEPG stripper off-gas while absorbing only a 

small percentage of the entering CO2. This creates a low-sulfur absorber off-gas that is recycled 

to the fuel combustor as low-energy-value fuel. The H2S-rich acid gas from the amine 

regenerator is sent to the methanol/H2S absorption system (M-404). 

Methanol/H2S absorption system (M-404) – This system utilizes a slip-stream of methanol from 

the alcohol separation area to absorb and recycle a controlled amount of H2S to the alcohol 

synthesis reactor (R-410) to maintain an appropriate H2S partial pressure for the metal-sulfide 

catalyst. The off-gas from the methanol absorber is sent to the LO-CAT sulfur recovery system 

(M-405). 

LO-CAT sulfur recovery system (M-405) – This system utilizes Merichem’s LO-CAT 

technology to remove H2S from the methanol absorber off-gas. LO-CAT is an iron-chelate-based 

process consisting of a venturi pre-contactor, liquid-filled absorber, air-blown oxidizer, air 

blower, solution circulation pump, and solution cooler. Elemental sulfur is produced in the 

oxidizer and stockpiled for disposal or for sale as an unconditioned product. The LO-CAT 

process was modeled to remove the H2S to a concentration of 10 ppmv in the CO2 vent. 

The turbo-expander for the recycle purge gas is designed as two stages with intermediate reheat 

in order to maximize power recovery. The inlets to both stages operate at 950°F (510°C). The 

pressure at the inlet of the first expander is dropped from more than 2,900 psi to 1,500 psi, with 

the assumption that it will be difficult to find expander technologies that operate at 2,900 psi. 

This maximum inlet pressure of 1,500 psi was derived from the data in Figure 11, which 

correlate maximum operating pressure as a function of inlet temperatures. The maximum 

operating pressure curve is based on ASTM pressure rating data and an operating envelope data 

point from Dresser-Rand. 

ASTM Maximum Pressure Rating Maximum Operating Pressure Dresser-Rand Data Point 

Pressure (PSIA) 

3500 

3000 

2500 

2000 

1500 

1000 

500 

0 

100 200 300 400 500 600 700 800 900 1000 1100 

Operating Temperature (Deg F)
 

Figure 11. Maximum operating pressure for syngas turbo-expander
 

32
 
 

 

 

 

 

 

 

 

 

 

 

 

The first turbine drops the pressure from 1,500 psi to 250 psi and the second turbine drops the 

pressure from 248 psi to 34 psi. An efficiency of 75% was assumed for the expanders. Although 

vendor quotations were requested for this equipment, vendors were reluctant to provide 

quotations due to the custom nature of the expanders and because vendors were generally 

unenthusiastic about spending the time to develop an estimate for which there is low likelihood 

of equipment purchase. Instead, literature sources, steam turbine information, Aspen IPE 

estimates, and data from previous projects done by Harris Group Inc. on steam turbines were 

used as a guideline to estimate the cost of the expanders. It is expected that the capital cost 

shown in Appendix B is conservative, as it is 3.5 times the price per kW of the quoted steam 

turbines in area 600. 

The recycle compressor is a low horsepower (706 hp in the base case) unit and was included in 

the quotation for the main syngas compressor. 

The reactor effluent is cooled to 110°F (43°C) through a series of heat exchangers and is 

maintained at high pressure. The effluent is first heat-exchanged with cooler process streams, 

lowering the temperature to 175°F (79.4°C). This is followed by air-cooled exchangers that 

reduce the temperature to 140°F (60°C). Finally, cooling water is used to drop the temperature to 

110°F (43°C). A knock-out drum (S-501) is used to separate the liquids (primarily alcohols) 

from non-condensable gas, comprised of unconverted syngas, CO2, and methane. The nonrandom two-liquid (NRTL) physical property package in Aspen Plus was used to model this 

flash operation as well as the alcohol condensation exchanger train. 

3.4.2 Area 400 Design Basis 

The design basis for the alcohol synthesis reactor system is presented in Table 13. The alcohol 

productivity targets for 2012 assume a 20% increase in specific activity over the 2010 State of 

Technology. This activity improvement was simulated through an activity factor adjustment in 

the mixed alcohol synthesis kinetic model. Design basis parameters for the other parts of the 

alcohol synthesis process area are presented in Table 14. 

33
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 13. Target and Design Basis Parameters for Alcohol Synthesis Reactor 

Technical Target Parameter 

Pressure, psi (bar) 3,000 (207) 

Single-pass CO conversion 29% 

Overall CO conversion 79% 

CO selectivity to alcohols (CO2-free) 81% 

CO selectivity to ethanol 63% 

Total alcohol productivity, g/kg-cat/h 368 

Ethanol productivity, g/kg-cat/h 160 

Design Basis Process Parameters Inlet Outlet 

Reactor temperature, °F (°C) 595 (313) 611 (322) 

Reactor pressure, psi (bar) 2,995 (206.5) 2,953 (203.6) 

Steam (shell side) temperature, °F (°C) 581 (305) 

Inlet gas composition 

H2:CO molar ratio 1.5 

CO2, mole % 14.0 

CH4, mole% 7.7 

Methanol, mole % 2.6 

H2S, ppmv 70 

H2O, wt % 0.04 

Reactor Design Parameters 

Reactor tube length, ft 60 

Total number of tubes 9,579 

Number of reactors 2 

Gas hourly space velocity, hour-1 5,000 

Catalyst particle size (cylindrical) 5 mm dia. x 5 mm length 

Reactor pressure drop, psi 42.2 

Tube outside diameter, inches 1.5 

Tube wall thickness, inches 0.16 

Reactor tube material A213 T22 (21⁄4 Cr, 1 Mo) 

34
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 14. Design Parameters for Compression, Power Recovery, and H2S/CO2 Systems 

Syngas Compressor Design Parameters 

Service Syngas + effluent recycle 

Inlet pressure 22.7 psia 

Outlet pressure 3,000 psia 

Inlet mass flow 350,450 lb/h 

Inlet molecular weight 17.2 lb/lb-mole 

Drive power source Electric motors 

Inter-stage cooling Process, air, cooling water 

Number of stages 6 

DEPG Acid Gas Removal (M-401) Design Parameters 

Absorber inlet pressure 2,937 psia 

Vapor feed rate 49,121 lb-mole/h 

H2S concentration 59 ppmv 

CO2 concentration 19.1 mole % 

Target CO2 removal from synthesis recycle loop 22.8% of inlet CO2 removed 

Solvent circulation rate 1,360 gpm 

Required refrigeration duty 23.1 MMBtu/h 

Stripping steam duty 13.9 MMBtu/h 

Cooling water duty 5.6 MMBtu/h 

Electrical consumption 3,310 kW 

Ammonia Absorption Refrigeration (M-402) Design Parameters* 

Required solvent temperature 30°F 

Refrigeration duty 23.4 MMBtu/h 

Cooling water 68.8 MMBtu/h 

Steam duty 40.6 MMBtu/h 

Electrical consumption 179 kW 

Amine Acid Gas Enrichment (M-403) Design Parameters 

Inlet molar flow rate 1,453 lb-mole/h 

H2S concentration 0.18 mole % 

CO2 concentration 87.2 mole % 

Selective amine concentration in solution 50 wt % 

Amine solution circulation rate 80 gpm required 

200 gpm maximum 

Treated gas H2S specification < 50 ppmv 

Steam duty 5.0 MMBtu/h 

Cooling water duty 5.0 MMBtu/h 

Electrical consumption 9 kW 

* Values for ammonia absorption refrigeration (M-402) duties were extrapolated from vendor quotation. 

35
 
 

 

 

 

 

 

 

 

 

 

 

Methanol/H2S Absorption (M-404) Design Parameters 

Inlet vapor H2S concentration 6.0 mole % 

Base case H2S recycle 1.38 lb-mole/h 

Maximum H2S recycle capacity 20 lb-mole/h 

Base case methanol rate 230 lb-mole/h 

Maximum methanol rate 2,000 lb-mole/h 

Base case refrigeration duty 0.26 MMBtu/h 

Maximum refrigeration duty 2.5 MMBtu/h 

Base case electrical consumption 71 kW 

Maximum electrical consumption 500 kW 

LO-CAT Sulfur Recovery (M-405) Design Parameters 

Sulfur recovery rate 936 lb/day
 

CO2 vent H2S concentration 10 ppmv
 

Expander/Power Recovery Design Parameters 

1st stage inlet pressure 1,500 psia 

1st stage inlet temperature 950°F 

1st stage power recovery 4.73 MW 

2nd stage inlet pressure 245 psia 

2nd stage inlet temperature 950°F 

2nd stage power recovery 6.74 MW 

36
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.4.3 Area 400 Equipment Cost Estimations 

The estimated purchased equipment costs and total installed costs for the alcohol synthesis area 

were gathered from various sources including technology licensors, industrial suppliers, 

published literature, and Aspen Icarus Process Evaluator 2006.5 computer software (Q1 2007 

cost basis). The sources of the cost data are summarized in Table 15. 

Table 15. Sources of Equipment Costs for Alcohol Synthesis Area 

Equipment or System Source of Equipment Costs 

Syngas compressors and motors Elliott – Ebara Group 

Compressor inter-stage air coolers GEA Rainey Corporation 

Compressor inter-stage KO drums Aspen Icarus Process Evaluator 2006.5 

Compressor inter-stage water coolers Aspen Icarus Process Evaluator 2006.5 

Alcohol synthesis reactors Hooper Engineered Vessels International 

Expanders and generators Aspen Icarus Process Evaluator 2006.5 

DEPG acid gas removal system Docter et al. [34] 

Ammonia absorption refrigeration system Colibri – BV 

Amine acid gas enrichment system Gary and Handwerk [35] 

Aspen Icarus Process Evaluator 2006.5 

Multiple selective amine technology licensors 

Methanol/H2S absorption system Aspen Icarus Process Evaluator 2006.5 

LO-CAT sulfur recovery system Adapted from a Merichem Company 

quotation for another project 

Table 16 presents the equipment list and cost estimates for the alcohol synthesis area of the plant. 

Table 16. Cost Estimate for Alcohol Synthesis Area 

Equipment Description TPEC TIC 

(2007 $k) (2007 $k) 

Synthesis gas compressors 32,537 59,057 

Alcohol synthesis reactor and effluent cooling/flash 11,851 29,551 

Purge gas expanders 9,961 17,929 

H2S/CO2 removal and sulfur recovery systems 13,586 28,199 

Heat integration 6,788 15,805 

Area A400 subtotal 74,722 150,541 

Additional equipment lists and cost breakdowns for the DEPG acid gas removal system, amine 

acid gas enrichment system, methanol/H2S absorption system, and LO-CAT sulfur recovery 

system are provided in Appendix H. 

3.5 Area 500: Alcohol Separation 

The following section presents an overview, basis for design, and cost estimates for construction 

of the alcohol separation area. 

37 
 

 

 

 

 

 

 

 

 

 

 

 

3.5.1 Area 500 Overview 

This area of the process recovers the alcohol products from the synthesis area and recycles 

unconverted material for improved overall conversion. The following steps take place in the 

alcohol separation area: 

• Molecular sieve for water removal from alcohol synthesis products 

• Two-step distillation section for recovery of ethanol and higher alcohol products. 

Figure 12 presents a simplified process flow diagram for the alcohol separation area of the plant. Flash 

From MAS 

Reactor Effluent 

Product Cooling 

Unconverted 

Syngas to Tar 

Reformer 

Molecular Sieve

(In Service)

Molecular Sieve

(Regeneration) 

Flash

Crude Alcohol Distillation Column

Methanol / Ethanol Splitter Column 

Methanol / 

Water Recycle 

to Gasifier 

Higher 

Alcohols 

to Storage 

Ethanol 

Fuel to 

Storage 

Recycle to 

Alcohol 

Synthesis 

Reactor 

Figure 12. Simplified process flow diagram of alcohol separation area 

The mixed alcohol stream from area 400 is sent to area 500 where it is de-gassed, dried, and 

separated into three streams: methanol, ethanol, and mixed higher-molecular-weight alcohols. 

Fifteen percent of the methanol stream is used to back-flush the molecular sieve drying column. 

This was deemed adequate based on available data for such mol-sieve purge operations [25]. The 

water/methanol stream from the purge is vaporized and fed to the gasifier (area 200) to satisfy 

part of its fluidization requirement and to recover carbon from the methanol. The remainder of 

the methanol is recycled back to the alcohol synthesis reactor (area 400) because methanol 

recycle has a positive impact on ethanol yields for The Dow Chemical Company’s catalyst. A 

slip-stream of this recycle is chilled to 30°F in order to dissolve H2S in the recycle and meet the 

70-ppm H2S specification at the alcohol synthesis reactor inlet. The ethanol and mixed alcohol 

streams are cooled and sent to product storage tanks. 

After the effluent from the alcohol synthesis reactor is cooled, the unconverted syngas and noncondensables are separated from the condensed alcohols in S-501. Although the majority of the 

lighter molecules leaving the synthesis reactor are separated in S-501, the condensed alcohol 

stream retains substantial quantities of H2, CH4, CO, and CO2 due to solubility at the high 

alcohol synthesis operating pressure. The condensed alcohol stream is then reduced in pressure 

from 2,937 psi to 60 psi and flashed in S-502 to reject the dissolved gases. The vapor from S-502 

38
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

is recycled to the tar reformer (area 300). The liquid from S-502 is reheated to the vapor phase, 

superheated, and sent to the molecular sieve dehydrator (S-503) for water removal. 

The molecular sieve dehydrator is based on previous ethanol studies [2, 22] and is assumed to 

have similar performance with mixed alcohols. The system consists of two parallel vessels 

containing fixed beds of porous alumino-silicate pellets, which are designed to selectively adsorb 

water. One molecular sieve vessel is in alcohol drying service while the material in the other 

vessel is being regenerated by recycled methanol. Once the material in drying service reaches 

water adsorption capacity, it enters the regeneration cycle and the regenerated vessel is placed in 

drying service. The methanol/water mixture resulting from the molecular sieve regeneration is 

cooled to 140°F (60°C) using a forced-air heat exchanger and separated from the uncondensed 

vapor in a flash drum. The vapor stream is recycled to the tar reformer and the water/methanol 

condensate is sent to the gasifier. Alumino-silicate molecular sieve products designed 

specifically for ethanol drying are commercially available through companies such as Grace [36] 

and UOP [37]. The capital cost for the unit was updated from recent vendor quotes. 

In the biochemical ethanol cases, the molecular sieve is used to dry ethanol after it is distilled to 

near the azeotropic concentration of ethanol and water (92.5 wt % ethanol). The adsorbed water 

is flushed from the molecular sieves with a portion of the dried ethanol and recycled to the 

rectification column. The water is ultimately rejected from the bottom of the distillation column. 

In a thermochemical process, where water concentration is low in the crude alcohol, it has been 

determined that drying the entire mixed alcohol stream prior to distillation operations is 

preferred. The adsorbed water is desorbed from the molecular sieves through a combination of 

depressurization and flushing with methanol. The estimated power consumption for the unit is 

factored into the Aspen Plus model and included as a plant operating cost. 

The dry mixed alcohol stream leaving the molecular sieve enters the first of two distillation 

columns, the crude alcohol column (D-504). The crude alcohol separation takes place in a typical 

distillation design incorporating a trayed column (D-504), an overhead condenser (H-504C), and 

a reboiler (H-504R). Essentially all of the methanol and 99% of the ethanol entering the column 

are recovered in the overhead product. The bottoms product consists of 1% of the feed ethanol, 

99% of the feed propanol, and all of the higher alcohols (butanol and heavier). The D-504 

bottoms stream is considered a co-product of the plant with a sale value of approximately 90% of 

its heating value (refer to Section 4.4 for details). The co-product is cooled and sent to product 

storage. The methanol/ethanol overhead from D-504 continues to a second distillation column, 

the methanol column (D-505), for further processing. 

The methanol/ethanol separation takes place in another typical distillation design, which includes 

a trayed column (D-505), an overhead condenser (H-505C), and a reboiler (H-505R). The 

ethanol is recovered from the bottoms of D-505 and contains 99% of the feed ethanol. The 

ethanol product contains less than 0.5 vol % methanol, meeting ASTM product specifications for 

fuel ethanol (refer to fuel-grade ethanol specifications in Appendix J). The ethanol is cooled and 

sent to product storage. The methanol stream is recycled to several destinations throughout the 

plant, as explained, and is not considered a saleable product. 

39
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.5.2 Area 500 Design Basis 

Table 17 summarizes the design basis for the equipment in the alcohol separation area, including 

the molecular sieve dryer system and alcohol distillation systems. 

Table 17. Design Basis Parameters for Alcohol Separation Area 

Molecular Sieve Design Parameters 

Inlet mass flow 

Inlet water content 

Target outlet water content 

Crude Alcohol Distillation Design Parameters 

129,421 lb/h 

5.6 wt % 

0.5 wt % 

Ethanol recovery in overhead 

Propanol recovery in bottoms 

Column diameter 

Number of theoretical trays 

Tray efficiency 

Number of actual trays 

Type of trays 

Tower and tray material 

99% 

99% 

9.0 ft 

60 

60% 

100 

Valve 

304SS 

Methanol / Ethanol Distillation Design Parameters 

Ethanol recovery in bottoms 

Methanol in bottoms ethanol product 

Column diameter 

Number of theoretical trays 

Tray efficiency 

Number of actual trays 

Type of trays 

Tower and tray material 

Heat Exchanger and Vessel Design Parameters 

99% 

0.5% 

12.0 ft 

81 

60% 

135 

Valve 

304SS 

Column overhead condenser cooling media 

Column reboiler heating media 

Column overhead accumulator sizing 10-m

Product rundown cooling media Pro

Air 

Low pressure steam 

inute liquid inventory 

3 vapor space factor 

cess + cooling water 

40
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.5.3 Area 500 Equipment Cost Estimations 

Table 18 presents the equipment list and cost estimates for the alcohol separation area of the 

plant. Costs for the molecular sieve drying system were extrapolated from a quotation by Delta-T 

Corporation provided for the biochemical ethanol design report [38]. All other equipment was 

priced using Aspen IPE 2006.5 software. 

Table 18. Cost Estimate for Alcohol Separation Area 

Equipment Description TPEC TIC 

(2007 $k) (2007 $k) 

Molecular sieve separator system 4,081 7,500 

Crude alcohol distillation 2,515 3,539 

Methanol/ethanol distillation 6,134 7,942 

Heat integration 493 1,326 

Area A500 subtotal 13,223 20,307 

3.6 Area 600: Steam System and Power Generation 

The following section presents an overview, basis for design, and cost estimates for construction 

of the steam system and power generation facilities. 

3.6.1 Area 600 Overview 

The present design includes a steam cycle that produces steam through recovery of heat from the 

hot process streams throughout the plant. Operations requiring steam include gasification, 

reforming, acid gas removal, acid gas enrichment, ammonia-water refrigeration system, and 

alcohol distillation. Steam is directly injected into the gasifier and the reformer. In the base case 

design, a target steam-to-carbon ratio of 2 can be maintained in the reformer with minimal 

additional steam feed. Steam is also used for indirect heat exchange in the distillation reboilers 

and acid-gas cleanup systems. All indirect heating applications include condensate return loops. 

Power for plant operations is produced from the steam cycle using a two-stage steam turbine 

with intermediate reheat to increase electricity production. The steam turbine efficiencies are 

assumed to be 75% and the generator mechanical efficiencies are assumed to be 97%. The first 

stage (M-602A) turbine drops the steam pressure from 1,321 psia to 350 psia. The second stage 

(M-602B) turbine drops the pressure from 345 psia to 3.5 psia. The steam inlets to both turbine 

stages are preheated to 1,000°F. Slip-streams of low pressure steam are extracted at 65 psia from 

the second stage for the uses described above. Power is also produced from the process turboexpander system in area 400 (K-412A,B). The plant energy balance is controlled by the rate of 

syngas combustion in the reformer catalyst regenerator. The energy from syngas combustion is 

recovered from the flue gas and contributes to the conversion to electrical energy in the steam 

turbine generators. The rate of syngas combustion is controlled to generate only enough 

electricity to supply the plant, minimizing electricity exported to the grid. The steam and power 

generation systems are shown in PFD-P820-A601 through PFD-P820-A604 in Appendix E. 

Low pressure (3.5 psia) steam exhaust from the turbine is cooled by exchanging heat with other 

process streams and is condensed in an air-cooled condenser (H-601). While an air-cooled 

41
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

exchanger is considerably more expensive than a water cooler, it has the advantage of reducing 

water consumption through evaporation in the cooling towers and was preferred in this design 

for environmental considerations. A condensate collection tank (T-601) gathers condensate 

returns from indirect heat exchange along with the steam turbine condensate. The condensate 

from the collection tank is treated in the hot condensate polishing unit (M-601C), mixed with 

treated makeup water, and sent to the condensate surge tank (T-602). From the surge tank, the 

water is pumped, heated to the saturation temperature, and sent to the deaerator (T-603) to 

remove any dissolved gases from the water. The water from the deaerator is pumped to a 

pressure of 1,351 psia and then pre-heated to its saturation (bubble point) temperature using a 

series of exchangers. Steam is generated by heat exchange with flue gas (H-611) and on the shell 

side of the alcohol synthesis reactor (R-410). This saturated steam is collected in the steam drum 

(T-605). To prevent buildup of solids, water is periodically discharged from the steam drum (also 

known as blowdown). Saturated steam from the steam drum is superheated with another series of 

exchangers to 1,000°F (538°C) before it goes to the steam turbine. 

The integration of the individual heat exchangers can be seen in the PFDs in Appendix E. The 

Aspen Plus model increases or decreases the boiler feed water flow rate through the steam cycle 

in order to close the heat balance. The control system in the plant will function in a similar 

fashion by controlling the flow rate of boiler feed water makeup to the steam system in order to 

maintain levels in the deaerator (T-603) and steam drum (T-605). 

All compressors, pumps, fans, etc. utilize electric motor drives and are modeled and quoted 

accordingly. An allowance of 1 MW of excess power is specified to account for miscellaneous 

usage and general electric needs (lights, computers, etc.). 

3.6.2 Area 600 Design Basis 

Table 19 presents a summary of the power requirements for each area of the plant. As all 

necessary power is generated in the plant, these power requirements define the basis for design 

of the steam system and power generation facilities. 

Table 19. Plant Power Requirements 

Plant Area Generated (kW) Consumed (kW) 

Feed handling and dryinga 0 

Gasification 7,112 

Tar reforming and gas cleanup 1,395 

Mixed alcohol synthesis, recycle compressor, 

turbo-expander, syngas compressor 11,467 51,724 

Alcohol separation and purification 808 

Steam system and power generation 52,944 1,519 

Cooling water and other utilities 798 

Miscellaneous 1,000 

Balance to grid 55 

Plant totals 64,411 64,411 

a Power costs for all feed handling, and for drying outside the plant, are included in the feedstock cost provided by 

INL. Energy for drying from 30% to 10% inside the plant is provided by process heat from flue gas. 

42 
 

 

 

 

 

 

 

 

 

 

 

Plant power demands and power production rates are designed to be balanced. Therefore, no 

significant excess power is claimed for by-product credits (sale to the grid) or bought from the 

grid. The plant design was conducted with an energy self-sufficiency requirement, and this was 

accomplished by burning portions of two syngas streams – 4% of the fresh, unreformed syngas 

and 25% of the unconverted syngas purged from the alcohol synthesis reactor recycle loop – in 

the tar reformer fuel combustor (area 300). While this does have a negative impact on the overall 

alcohol yields of the process, it negates the purchase of natural gas or grid power during normal 

operations. 

3.6.3 Area 600 Equipment Cost Estimations 

Table 20 presents the equipment list and cost estimates for the steam system and power 

generation area of the plant. 

Table 20. Cost Estimate for Steam System and Power Generation Area 

Equipment Description TPEC TIC 

(2007 $k) (2007 $k) 

Water treatment system (RO and EDI units) 2,465 2,835 

Boiler feed water/steam/condensate handling 1,333 3,226 

Extraction steam turbines/generators 13,272 23,890 

Steam turbine condenser 6,215 8,701 

Heat integration 2,560 7,193 

Area A600 subtotal 25,845 45,845 

3.7 Area 700: Cooling Water and Other Utilities 

The following section presents an overview, basis for design, and cost estimates for construction 

of the cooling water facilities and other plant utility equipment. 

3.7.1 Area 700 Overview 

The cooling water and other utilities systems are shown in PFD-P820-A701 and PFD-P820­

A702 in Appendix E. A mechanical draft cooling tower (M-701) provides cooling water to 

several heat exchangers in the plant. The tower utilizes fans to force air through circulated water. 

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

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

instrument air is delivered at a pressure of 115 psia and a moisture dew point of -40°F (-40°C), 

and it is oil free. 

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

• Firewater storage tank (T-702) and firewater pump (P-702) for fire emergencies 

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

• Olivine truck dump with scale (M-702) and olivine/MgO loading system (T-706) 

• Tar reformer catalyst loading system (T-707) 

43
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

•	 Ammonia storage tank (T-704), ammonia pump (P-704), BFW chemical storage tank (T708), and BFW chemical pump (P-707) for boiler feed water treatment. 

•	 Caustic (NaOH) storage tank (T-705) and pump (P-705). 

3.7.2 Area 700 Design Basis 

Table 21 summarizes the design basis parameters for the plant’s cooling water system. 

Table 21. Design Basis for Cooling Water System 

Cooling Water System Design Parameters 

Cooling water supply temperature (TS) 90°F 

Cooling water return temperature (TR) 110°F 

Cooling water supply rate (FS) 11,720 gpm 

Losses from evaporation (FEVAP) 0.00085 * FS * (TR - TS) 

Losses from drift 0.002 * FS 

Cycles of concentration (C) 5 

Required blowdown FEVAP / (C – 1) 

Cooling water makeup pressure 14.7 psia 

Cooling water makeup temperature 60°F 

Water losses from the cooling water system include evaporation, drift (water entrained in the 

cooling tower exhaust air), and tower basin blowdown. Drift losses were estimated to be 0.2% of 

the cooling water supply rate. Evaporation losses and blowdown were calculated based on 

information and equations in Perry et al. [17] (as shown in Table 21). 

3.7.3 Area 700 Cost Estimation 

Table 22 presents the equipment list and cost estimates for the cooling water and other utilities 

area of the plant. 

Table 22. Cost Estimate for Cooling Water and Other Utilities Area 

Equipment Description TPEC TIC 

(2007 $k) (2007 $k) 

Plant/instrument air systems 307 583 

Cooling water system 660 1,480 

Flue gas scrubber and stack 836 1,918 

Purge water treatment 361 913 

Catalyst, chemicals, and product storage 2,120 4,160 

Miscellaneous utility systems 276 507 

Area A700 subtotal 4,560 9,560 

44
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.8 Additional Process Design Information 

Table 23 contains additional information used in the Aspen Plus model and design basis. 

Table 23. Utility and Miscellaneous Design Information 

Item Design Information 

Ambient air conditions a,b,c 

Pressure: 14.7 psia 

Tdry bulb: 90°F 

Twet bulb: 80°F 

Air composition (mol %) 

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

: 

Pressure drop allowance 

Syngas compressor intercoolers = 2 psi per exchanger 

Air pre-heat exchangers = 1 psi 

Heat exchangers = 5 psi 

a In the GPSA Engineering Data Book [39], see Table 11.4 for typical design values for dry bulb and wet bulb 

temperature by geography. Selected values would cover summertime conditions for most lower 48 states.

b In Weast [40], see F-172 for composition of dry air. Nitrogen value adjusted slightly to force mole fraction closure 

using only N2, O2, Ar, and CO2 as air components.
 c In Perry et al. [17], see psychrometric chart, Figure 12-2, for moisture content of air.
 

3.9 Pinch Analysis 

A pinch analysis was performed to design the heat exchange network of the biomass gasification 

to ethanol production process. The concept of pinch technology offers a systematic approach for 

optimizing energy integration in process design. Temperature and heat flow (Q) data were 

gathered for the hot process streams (streams cooled in the process), cold process streams 

(streams heated in the process), and utility streams like steam, flue gas, and cooling water. A 

composite curve, a graphical presentation of stream temperature and heat flow (Q) data, was 

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

at the pinch point. From this shifted graph, a grand composite curve is constructed that plots the 

difference in heat flow between the hot and cold composite curves as a function of temperature. 

The total hot duty equals the total cold duty, as the Aspen Plus model is designed to adjust the 

water flow rate through the steam cycle until the heat balance in the system is closed. The 

minimum vertical distance between the curves is ΔTmin, which is theoretically the smallest 

approach temperature in the exchanger network. For this design, the pinch occurs at ~ Q = 285 

MMBtu/h where the upper and lower pinch temperatures are 331.6°F and 316.3°F, respectively, 

resulting in ΔTmin of 15.3°F. 

There are a multitude of options for heat integration network design, depending on objectives 

and available capital. In general, as efficiency is increased in an exchanger network design, the 

required heat transfer area increases and therefore capital cost increases. The objective of this 

network design is to lower energy consumption and boost product yields, keeping in mind that 

very low ΔT can lead to impractically large exchangers and high capital costs. Figure 13 shows 

the cumulative hot and cold pinch curves after they were simplified by removing dedicated steam 

generation (from the alcohol synthesis reaction), indirect steam use systems (distillation reboilers 

and acid gas removal), and water- and air-cooling utilities. A specific network was designed 

based on this philosophy and the heating and cooling curves shown in Figure 13. The exchangers 

are listed in the equipment list in Appendix B, with the scaling stream entries listed as “Pinch.” 

45
 
 

 

 

 

 

 

 

The exchangers in the heat integration network are also shown on the process flow diagrams in 

Appendix E, with details of each exchanger listed in Appendix M. A total of 53 exchangers were 

included in the design with a total installed cost of approximately $27MM. Aspen Icarus Process 

Evaluator is the source of the equipment cost and installation estimates for the heat integration 

network. Temperature (°F) 

2000 

1800 

1600 

1400 

1200 

1000 

800 

600 

400 

200 

0 

Hot Process 

Cold Process 

0 100 200 300 400 500 600 700 800 900 1,000 

Heat Flow (MMBtu/Hr) 

Figure 13. Pinch analysis composite curve 

3.10 Energy Balance 

Energy integration is important to the overall economics and efficiency of the process. Therefore 

a detailed understanding of how and where the energy is utilized and recovered is required. 

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

Plus simulation. Comparing the process energy inputs and outputs allows the energy efficiency 

of the process to be quantified. Tracing energy transfer between process areas also makes it 

possible to identify potential improvements to the energy efficiency. 

This design for thermochemical conversion of cellulosic biomass was constructed with the goal 

of energy self-sufficiency. Natural gas inputs that could be used to fire the char combustor and 

fuel combustor have been eliminated. Instead, a slip-stream of unreformed syngas and a portion 

of the unreacted syngas from the alcohol synthesis reactor are used to meet the fuel demand. The 

downside to this is a decrease in ethanol yield. Because no electricity is purchased from the grid, 

the integrated combined heat and power system must supply all steam and electricity needed by 

the plant. A negligible amount of electricity is sold as a co-product. The only saleable products 

are assumed to be fuel ethanol and higher-molecular-weight mixed alcohol co-product. 

46
 
 

 

 

 

 

 

The distribution of the total energy content in dry wood (LHV basis) is shown in Figure 14. The 

total energy in the dry portion of the feed (0% moisture) is approximately 1,470 MMBtu/h. 

Figure 14 shows that approximately 45% of the LHV is recovered in ethanol and higher alcohol 

products. A significant amount (23%) is lost through air-cooled exchangers, and about 15% is 

lost because of moisture in the feed and other water input to the process. Electrical and thermal 

losses account for 11%. 

Further optimization of the heat integration network can be evaluated. However, it is our 

expectation that further heat integration for this process design will have marginal impact on the 

MESP and should therefore be addressed only if there is a detailed plant design. 

Misc (sand,ash,air 

in,sulfur etc.) 

2% 9% 6%

Figure 14. Overall energy analysis (dry wood LHV basis) 

Ethanol Product 

40% 

Higher Alcohol 

Product 

5% 

Water (Balance) Moisture in Wood 

Cooling Tower 

Evaporation 

Flue Gas 

1% 

Gasifier Loss 

2% 

Air Cooled 

Exchangers 

23% 

Electrical and 

Thermal Losses 

11% 

1% 

47
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

3.11 Sustainability Metrics 

An important aspect of evaluating biomass-derived fuel processes is the quantification of life 

cycle resource consumption and environmental emissions. Life cycle assessment (LCA) provides 

a framework from which the environmental sustainability of a given process may be quantified. 

This section presents some key LCA input data from the thermochemical ethanol model. These 

data are not an exhaustive life cycle inventory and are only intended to facilitate populating the 

direct resource consumption and environmental emissions attributed to producing ethanol via 

indirect gasification and mixed alcohol synthesis. 

Table 24 quantifies the consumption of resources and emissions from the plant. Two cases are 

presented in the table for water consumption values – the base design case and an optimized 

water consumption scenario recently developed by NREL and Harris Group Inc. The optimized 

scenario efficiently utilizes reverse osmosis (RO) and brine recovery units to recycle and reuse 

water in the plant. NREL continues to evaluate options for further reduction of water 

consumption and environmental emissions without significant economic penalty. Consumption 

and discharge of materials other than water are equal in both cases. 

Table 24. Sustainability Metrics for Thermochemical Ethanol 

Base Case Optimized Water Consumption 

(if varied from Base Case) 

Flow rate 

lb/h (gal/h) 

Per gal EtOH 

lb/gal (gal/gal) 

Flow rate 

lb/h (gal/h) 

Per gal EtOH 

lb/gal (gal/gal) 

Water inputs 

Water in biomass 78,736 (9,463) 10.23 (1.23) 

Water in air to char combustor 9,320 (1,120) 1.21 (0.15) 

Water in air to catalyst regenerator 4,950 (595) 0.64 (0.08) 

Water in air to LO-CAT 2 (0) 0.00 (0.00) 

Makeup water to cooling tower 91,076 (10,946) 11.84 (1.42) 24,070 (2,893) 3.13 (0.38) 

Makeup water to boiler system 76,348 (9,176) 9.92 (1.19) 101,798 (12,234) 13.23 (1.59) 

Water outputs 

Ash wetting 243 (29) 0.03 (0.00) 

Cooling tower drift 11,729 (1,410) 1.52 (0.18) 

Cooling tower evaporation 98,651 (11,856) 12.82 (1.54) 101,779 (12,232) 13.23 (1.59) 

Cooling tower blowdown to WWT 24,663 (2,964) 3.21 (0.39) 0 (0) 0 (0) 

Other wastewater to WWT 28,828 (3,465) 3.75 (0.45) 7,710 (927) 1.00 (0.12) 

Flue gas 128,420 (15,434) 16.69 (2.01) 129,821 (15,602) 16.87 (2.03) 

CO2 exhaust 34 (4) 0.00 (0.00) 

Syngas scrubber solids 52 (6) 0.01 (0.00) 

Total water inputs 260,432 (31,299) 33.85 (4.07) 218,876 (26,305) 28.45 (3.42) 

Total water outputs 292,620 (35,167) 38.03 (4.57) 251,370 (30,210) 32.67 (3.93) 

Note: Water output is larger than water input because water is produced in the process. 

48
 
 

 

 

 

 

 

 

Table 24 (continued). Sustainability Metrics for Thermochemical Ethanol 

Flow rate Per gal EtOH 

lb/h lb/gal 

Other inputs 

Biomass (water-free) 183,718 23.9 

Olivine 538 7.0 x 10-2 

Magnesium oxide (MgO) 7.0 9.1 x 10-4 

Tar reforming catalyst 12.0 1.6 x 10-3 

Alcohol synthesis catalyst 19.4 2.5 x 10-3 

Caustic 40.0 5.2 x 10-3 

Boiler chemicals 3.3 4.3 x 10-4 

Cooling tower chemicals 1.2 1.5 x 10-4 

Diesel fuel 70.9 9.2 x 10-3 

LO-CAT chemicals 3.2 4.1 x 10-4 

DEPG makeup 1.8 2.3 x 10-4 

Amine makeup 0.2 2.3 x 10-5 

Air demand 

Carbon dioxide (CO2) 360 4.7 x 10-2 

Nitrogen (N2) 541,603 70.4 

Oxygen (O2) 165,948 21.6 

Air emissions 

Carbon dioxide (CO2) 232,769 30.2 

Sulfur dioxide (H2S) 9.4 1.2 x 10-3 

Nitrogen dioxide (NO2)

a 8.6 x 10-3 1.1 x 10-6 

Nitrogen (N2) 541,810 70.4 

Oxygen (O2) 25,738 3.3 

Waste streams 

Olivine, MgO, catalyst, ash, sulfate 2,708 3.5 x 10-1 

Alcohol synthesis catalystb 23.3 3.0 x 10-3 

Products 

Ethanol 51,000 6.63 

Mixed higher alcohols 6,448 8.4 x 10-1 

Sulfur 39 5.1 x 10-3 

a Based on nitrogen in fuel. Additional NOx from N2 in combustion air is not quantified. b 20% additional mass applied to fixed bed catalyst waste to account for deposits. 

Water usage is an increasingly important aspect of plant design, specifically for ethanol plants. 

Most U.S. ethanol plants reside in the Midwest where water supply is a significant concern [41]. 

For several years, significant areas of water stress have been reported during the growing season. 

In addition, livestock and irrigation operations compete for the available resources. 

Table 25 presents a comparison of water consumption volumes for gasoline from crude oil, 

ethanol from corn, and ethanol from the thermochemical pathway. The water consumption 

49
 
 

 

 

 

 

 

 

 

 

values for gasoline and corn ethanol are from a report by Wu et al. from Argonne National 

Laboratory [42]. The conversions of ethanol to gasoline energy equivalents use a factor of 1.52, a 

result of dividing the energy content of gasoline by that of ethanol (115,500 Btu per gallon 

gasoline / 76,000 Btu per gallon ethanol). 

Table 25. Comparison of Water Usage for Transportation Fuel Processes 

Transportation Fuel Description Gal Water per 

Gal Product 

Gal Water per 

Gal Gasoline Equiv. 

Crude oil refining to gasoline 2.8–6.6 2.8–6.6 

Corn ethanol 10.0–323.6 15.2–491.9 

Thermochemical cellulosic ethanol 

(base case) 2.6 4.0 

Thermochemical cellulosic ethanol 

(optimized water consumption case) 2.0 3.0 

50
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

4 Process Economics 

The first steps of the economic analysis are to determine the total capital investment (TCI), 

variable operating costs, and fixed operating costs associated with constructing and operating the 

plant. With these estimated costs, a discounted cash flow analysis is then used to determine the 

production cost of ethanol for a project having a net present value of zero at an assumed internal 

rate of return (IRR). This section describes the cost areas and the necessary assumptions required 

to complete the discounted cash flow analysis. 

4.1 Total Capital Investment (TCI) 

The following discussion summarizes the purchased and installed equipment costs presented in 

Section 3 of the design report and describes how the total installed costs (TIC) for the plant serve 

as the basis for determining the total capital investment (TCI). 

As previously discussed, the capital costs for each plant area are based on data from various 

sources including technology licensors, industry equipment suppliers, cost estimates from Harris 

Group Inc., published literature, and Aspen Icarus Process Evaluator 2006.5 cost estimating 

software, which is based on first quarter 2007 costs. The purchased costs for the equipment and 

installation factors are used to determine the installed equipment cost. Table 26 presents a 

summary of total purchased equipment costs (TPEC), installation factors (finstallation), and total 

installed costs (TIC) for each area of the plant, the bases for which are discussed in further detail 

in Section 3, Appendix A, and Appendix B. 

Table 26. General Cost Factors in Determining Total Installed Equipment Costs 

Area Process Description TPEC f Installation TIC 

(MM$) (MM$) 

100 Feed handling and dryinga Capital costs included in feedstock price 

200 Gasification 18.7 2.31 43.3 

300 Gas cleanup 12.6 2.14 26.9 

400 Alcohol synthesis 74.7 2.01 150.5 

500 Alcohol separation 13.2 1.54 20.3 

600 Steam plant and power 25.8 1.77 45.8 

700 Cooling water and other utilities 4.6 2.10 9.6 

ISBL (Areas 100–500) 119.3 2.02 241.0 

OSBL (Areas 600–700) 30.4 1.82 55.4 

Total 149.7 1.98 296.4 

a Investment costs for feed handling and drying are included in the per-unit woody feedstock price. 

The indirect costs (non-manufacturing fixed-capital investment costs) are estimated using factors 

based on the total direct cost (TDC). The TDC is shown in Table 28. The factors are summarized 

in Table 27 as percentages of total purchased equipment cost (TPEC), total direct cost (TDC), 

and fixed capital investment (FCI), which is equal to the sum of TDC and total indirect costs. 

51 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 27. Cost Factors for Indirect Costs 

Indirect Costs % of TPEC % of TDCa % of FCIa 

Prorated expenses 20.4 10.0 6.3 

Home office and construction fees 40.9 20.0 12.5 

Field expenses 20.4 10.0 6.3 

Project contingency 20.4 10.0 6.3 

Other costs (start-up and permits) 20.4 10.0 6.3 

Total indirect costs 122.7 60.0 37.5 

Working capital 5.0 

a Percentages of TDC and FCI exclude land purchase cost. 

With the calculated total installed cost and assumptions for indirect costs and working capital, 

the fixed capital investment (FCI) and total capital investment (TCI) are calculated. Table 28 

presents a summary of these calculations. 

Table 28. Project Cost Worksheet 

Total purchased equipment cost (TPEC) $149,720,048 

Installation factor 1.980 

Total installed cost (TIC) $296,447,211 

Other direct costs 

Land (not depreciated) $1,610,000 

Site development 4.0% of ISBL $9,641,711 

Total direct costs (TDC) $307,698,921 

Indirect costs % of TDC (ex Land) 

Prorated expenses 10.0% $30,608,892 

Home office and construction fees 20.0% $61,217,784 

Field expenses 10.0% $30,608,892 

Project contingency 10.0% $30,608,892 

Other costs (start-up and permits) 10.0% $30,608,892 

Total indirect costs 60.0% $183,653,353 

Fixed capital investment (FCI) $491,352,274 

Working capital 5.0% of FCI (ex Land) $24,487,114 

Total capital investment (TCI) $515,839,388 

TCI / TPEC 3.445 

FCI Lang Factor = FCI / ISBL TPEC 4.118 

TCI Lang Factor = TCI / ISBL TPEC 4.323 

52
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

4.2 Variable Operating Costs 

There are two types of operating costs – variable and fixed. The following sections present 

assumptions and calculation approaches to determine the variable and fixed operating costs 

associated with thermochemical ethanol production. 

The basis for variable operating costs considered in this analysis is summarized in Table 29. A 

detailed summary of variable plant operating costs is presented in Appendix C. 

Table 29. Summary of Variable Operating Costs 

Variable Information and Operating Cost (cost year in parentheses) 

Gasifier bed 

material 

Synthetic olivine and MgO. Initial fill then a replacement rate of 0.01 wt % of 

circulation or 7.2 wt % per day of total inventory. Delivered to site by truck with selfcontained pneumatic unloading equipment. Disposal by landfill. 

Olivine price: $172.90/ton (2004) [43]. 

MgO price: $365/ton (2004) [44]. 

Tar reformer 

catalyst 

To determine the amount of catalyst inventory, the tar reformer was sized for a gas 

hourly space velocity (GHSV) of 2,476/h based on the operation of the tar reformer 

at NREL’s TCPDU. GHSV is measured at standard temperature and pressure. 

Initial fill then a replacement rate of 0.1 wt % of catalyst inventory per day. 

Price: $17.74/lb (2007) based on NREL calculations using metals pricing and costs 

for manufacturing processes. 

Alcohol synthesis Initial fill then replaced every 2 years based on expected catalyst lifetime. 

catalyst Catalyst inventory based on GHSV of 5,000/h.
 

Price: $29.70/lb (2007) based on NREL calculations using metals pricing and costs
 

for manufacturing processes.
 

Salvage value: $4.12/lb (2007) based on metals reclamation quotations.
 

Solids disposal Price: $36/tonne (1998) [16] 

Diesel fuel Usage: 10 gal/h plant-wide use. 

2012 price projection: $21.29/MMBtu (2009) [45] = $2.86/gal at 0.85 specific 

gravity [corrected to 2007 currency by EIA index (2007 = 106.2, 2009 = 109.8)]. 

Water makeup Price: $0.22/tonne (2001) [16] = $0.20/ton. 

Chemicals Boiler feed water chemicals – Price: $1.40/lb (1991) [2]. 

Cooling tower chemicals – Price: $1.00/lb (1999) [2]. 

Caustic – Price: $150/dry ton (2010) [38]. 

LO-CAT chemicals – Price: $408/ton sulfur produced (2009) from NREL/Harris 

Group Inc. estimates based on other projects.
 

DEPG makeup – Price: $56.55/million lb acid gas removed (2004) [46].
 

Selective amine makeup – Price: $16.94/million lb acid gas removed (2010).
 

Wastewater The wastewater is sent to on-site treatment facility.
 

Price: $0.53/tonne (2001) [16]
 

Note: Costs shown in Table 29 were updated to 2007 dollars using an inorganic cost index [47].
 

Values used: 1991 = 125.6; 1998 = 148.7; 1999 = 149.7; 2001 = 158.4; 2004 = 172.8; 2007 = 203.3; 2010 = 233.9.
 

53 
 

 

 

 

 

 

 

 

 

 

4.3 Fixed Operating Costs 

This section considers the plant’s fixed operating costs, which include employee salaries and 

benefits, overhead, plant maintenance costs, insurance, and taxes (other than income taxes). The 

fixed operating costs used in this analysis are shown in Table 30 (labor costs) and Table 31 

(other fixed costs). 

Table 30. Salary Costs for Plant Employees 

Position Title Salary 

(2007) 

Number of 

Positions 

Total Cost 

(2007) 

Plant manager $147,000 1 $147,000 

Plant engineer $70,000 1 $70,000 

Maintenance supervisor $57,000 1 $57,000 

Laboratory manager $56,000 1 $56,000 

Shift supervisor $48,000 5 $240,000 

Lab technician $40,000 2 $80,000 

Maintenance technician $40,000 16 $640,000 

Shift operators $48,000 20 $960,000 

Yard employees $28,000 12 $336,000 

Clerks and secretaries $36,000 3 $108,000 

Total salaries (2007) $2,694,000 

Note: Labor costs are indexed, if necessary, to values from the U.S. Bureau of Labor Statistics
 

(http://data.bls.gov/cgi-bin/srgate CEU3232500008)
 

Because the salaries listed above do not include benefits, a general overhead factor is applied. 

The factors for other fixed costs in Table 31 also include general plant maintenance, plant 

security, janitorial services, communications, etc. A detailed summary of fixed plant operating 

costs is presented in Appendix C. 

Table 31. Other Fixed Operating Costs 

Cost Item Factor Cost 

Benefits and general overhead 90.0% of total salaries $2,424,600 

Maintenance 3.0% of fixed capital investment (FCIa

) $14,694,771 

Insurance and taxes 0.7% of fixed capital investment (FCIa

) $3,428,780 

Total other fixed operating costs (2007) $20,548,151 

a Percentages of FCI exclude land purchase cost. 

54
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

4.4 Value of Higher Alcohol Co-Products 

The alcohol synthesis process will create higher-molecular-weight alcohols in addition to 

ethanol. The valuation of this co-product is a function of the market in which it can be sold. Two 

extreme cases are considered here. At the high end, the co-product might be sold into the 

chemical market, commanding a high value of approximately $4 per gallon [48]. However, it is 

unlikely that the market would support the co-products from more than a few biomass plants. 

Therefore, because this design report assumes nth-plant economics, the economic analysis does 

not take credit for higher alcohols sold into the chemical market. 

At the low end, the co-product could command a value as a heating or transportation fuel with 

minimal ASTM standards on its specifications. In this scenario, the co-product would be priced 

relative to other heating and transportation fuels, such as gasoline, based on the heating value. 

This co-product valuation approach is applied for the base case economic analysis and 

calculation of MESP. Table 32 presents the co-product pricing calculations, which are based on 

the projected 2012 wholesale value of gasoline from the U.S. Energy Information Administration 

[49], inflation-adjusted to reflect a 2007 pricing basis. The calculated value of the mixed alcohols 

co-product is $2.11 per gallon based on the energy value of the mixed alcohols stream. The coproduct value conservatively assumed for the technoeconomic analysis is $1.88 per gallon 

(approximately 90% of the calculated value). 

Table 32. Higher Alcohols Co-Product Pricing 

Description Value 

EIA projected 2012 wholesale gasoline price (U.S. 2007$) $2.62/gallon 

Gasoline energy value 5.218 MMBtu/bbl 

Volume conversion 42 gallons/bbl 

EIA projected 2012 wholesale gasoline value (U.S. 2007$) $21.09/MMBtu 

Higher alcohols product energy value 

(from Aspen Plus simulation results) 100,149 Btu/gallon 

Calculated higher alcohols product value (U.S. 2007$) $2.11/gallon 

Higher alcohols product value for analysis (U.S. 2007$) $1.88/gallon 

4.5 Minimum Ethanol Selling Price 

Once the capital and operating costs are determined, a minimum ethanol selling price (MESP) is 

determined using a discounted cash flow rate of return analysis. The general methodology used 

is identical to that applied by Aden et al. (2002) [22]. The MESP value represents the selling 

price of ethanol assuming a 30 year plant life and 40% equity financing with 10% internal rate of 

return and the remaining 60% debt financed at 8% interest. The base case economic parameters 

used in this analysis were presented in Section 1.3 and summarized in Table 1. The resulting 

base case MESP is $2.05 per gallon of ethanol in U.S. 2007$. Sensitivity analysis was performed 

to examine the variability of the MESP as a function of changes in various process parameters, 

raw material costs, and financial assumptions. The results of the sensitivity analysis are 

discussed in Section 5. 

55
 
 

 

 

 

 

 

 

 

 

 

 

5 Process Economics and Sensitivity Analyses 

The cost of ethanol as determined in the previous section was derived using technology that has 

been developed and demonstrated or is currently being developed as part of the U.S. Department 

of Energy’s biomass research programs, which includes collaboration with industrial partners. 

The combination of all process, market, and financial targets as well as other assumptions 

presented in this design report result in a minimum ethanol selling price of $2.05 per gallon. A 

summary of the costs contributing to the total ethanol selling price is presented in Figure 15 and 

tabulated in Appendix C. 

Figure 15. Cost contribution details from each process area 

This cost contribution chart shows co-product credits for higher alcohols from the alcohol 

synthesis area, for electricity from the syngas expansion area, and for electricity from the steam 

system and power generation area. However, the process was adjusted so that the electricity 

generated is balanced by the electricity required by all other areas, so there is no net credit for 

electricity generation. 

The total cost of ethanol production, as indicated by the MESP, is determined from the 

amalgamation of various assumptions, both technical and financial. The 2012 research targets for 

cost-competitive ethanol (Appendix I) were the basis for this design. Financial and market 

assumptions include future demand for products like higher alcohols, financing criteria like 

internal rate of return (IRR) for equity investors, and interest on debt. In addition, potential 

56
 
 

 

 

 

variability in equipment design, installation, and construction costs may impact economics. 

Although the assumptions and estimated plant capital costs are reasonable, it is important to 

consider the impact of deviations from base case values. Sensitivity analysis results for project 

parameters identified by subject matter experts as critical factors in determining an accurate 

MESP are discussed here. The results for the sensitivity analysis are summarized in Figure 16. 

The data are organized by potential impact of parameter deviation on MESP value (highest to 

lowest). Most of the sensitivity scenarios represent a deviation in a single project parameter (e.g., 

higher alcohol co-product value or plant life) with all other parameters remaining constant at 

base case values. One exception is the scenario in which plant feedstock is varied from woody 

biomass to corn stover, requiring the entire feed composition (contents of carbon, hydrogen, 

nitrogen, sulfur, oxygen, ash, and moisture) to change simultaneously. 

Note that each sensitivity scenario in Figure 16 has an associated deviation value from the base 

case. If a deviation bar is not visible, then the impact on MESP is negligible. All sensitivity 

scenarios are discussed further in the following sections. The case numbers in the text refer to the 

numbers shown in the labels of Figure 16. 

Figure 16. Results of sensitivity analyses 

57
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

5.1 Financial Scenarios 

Deviations from base case financial and market parameters can have significant impacts on the 

MESP. The base case financial assumptions are shown in Table 1. Case 1 in Figure 16 shows the 

possible savings realized by economies of scale from changes to the plant size. It is assumed in 

the plant size sensitivity scenario that the same technology is utilized for all plant sizes. A major 

constraint to larger plant sizes will be significantly higher feedstock transportation costs. An 

increase in required IRR for the project from 10% to 20% (case 2) would increase the MESP 

from $2.05 to $2.79 per gallon (an increase of 36%). An interest rate increase (case 9) from the 

base value of 8% to 12% increases MESP to $2.19 per gallon (an increase of 7%). Successful 

R&D and demonstration projects would reduce investment and lending risks and ease the 

expected rate of return to investors and minimize loan interest rates, reducing the MESP toward 

the base case value. 

Variability in total capital investment (TCI) can also have a major impact on MESP. Conceptual 

designs like this generally result in accuracy of ±10% to 30% for capital investment estimation 

[16]. Applying a variability range of –10% to +30% to a TCI sensitivity (case 5) results in an 

MESP range of $1.92 to $2.43 per gallon (–6% to +19%). The impact of installation factors (case 

4) yields nearly identical results because this study is based on factored estimates. 

5.2 Feedstock 

Although the plant design basis specifies woody biomass feedstock from southern pine, there is 

potential for other feedstocks to be introduced as feed, based on future biomass availability. The 

negative impacts of higher ash and moisture contents on the MESP are quantified in case 6 and 

case 8. Product yields decrease for higher moisture feedstocks because more energy is used for 

drying. The impacts of changes in the feedstock cost (case 7) and the use of a corn stover 

feedstock (case 3) – the feedstock for NREL’s biochemical process 2012 target base case design 

[38] – are also shown. The high ash, low carbon, and high oxygen content in corn stover results 

in a 20% increase in the MESP (case 3). 

5.3 Gasification and Gas Cleanup 

The impact of changes to the capital costs of the gasifier and reformer is shown in case 17. 

Additional heat losses from the gasifier (case 25) will decrease yields, leading to increases in the 

MESP. 

Increasing methane conversion (case 21) to 90% in the reformer from a base case value of 80% 

will decrease the MESP by 2%. There is a potential for realizing this savings by the 2012 

demonstration because NREL’s reforming catalyst has shown such conversions under some 

conditions. The impacts of changes to the reforming catalyst cost (case 28) and longevity (case 

15) are also shown. 

5.4 Alcohol Synthesis 

Case 14 shows that improving the alcohol synthesis catalyst activity by 17% can reduce the 

MESP by 2% (keeping all other performance factors constant; the economic benefit can be 

slightly higher if the process is further optimized to take full advantage of the improved activity). 

Case 10 shows that the compressor capital cost and energy savings from operating the synthesis 

reactor at 1,500 psi (compared to 3,000 psi in the base design) can reduce the MESP by 6%, 

assuming there are no other changes to the process. Extending the catalyst lifetime (case 19) 

58
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

from the base case of 2 years and reducing the catalyst cost from the base case value (case 18) 

will also reduce the MESP. 

5.5 Co-Product Value 

The higher alcohol co-product value for the base case is calculated from the product’s heating 

value and the 2012 EIA projected gasoline value (Table 32). In the base case the co-product 

value is about 90% of the gasoline equivalent higher heating value. It is assumed that there will 

be demand for the higher-molecular-weight mixed alcohols in the liquid fuels market. This 

sensitivity analysis (case 22) shows the MESP impact from higher and lower co-product values, 

which can represent variability in demand. A co-product valuation range of 111% to 67% of the 

gasoline equivalent cost by heating value (125% to 75% of the heating value for the base case) 

results in an MESP range of $1.99 to $2.11 per gallon (–3% to +3%). 

6 Conclusions 

This analysis shows that biomass-derived ethanol from a thermochemical conversion process 

continues to show the potential to be cost-competitive with gasoline. The gasoline equivalent 

price by higher heating value is $3.11 per gallon, not including retail taxes, tax credits, or costs 

for distribution, blending, and marketing. The sensitivity analysis shows there is less than ±20% 

prediction uncertainty in MESP resulting from a ±30% uncertainty in the total capital cost 

estimation. While there is always a chance of large escalations in capital costs, acquisition of 

recent estimates and vendor quotes for most of the equipment reduces the probability of gross 

over- or under-estimation of costs. All processes used in the plant design are commercially 

available with the exception of the gasifier and tar reformer. Experimental results continue to 

approach the 2012 technical targets (details in Appendix I), meaning that the predictions 

presented in this report are closer to practice than those of the previous report published in 2007. 

Bench-scale experiments have shown that, in areas like methane conversion, actual performance 

can surpass the technical targets, resulting in reduced actual costs. 

7 Future Work 

The ultimate success of OBP’s program to demonstrate technologies capable of producing costcompetitive ethanol via biomass gasification by 2012 will depend on final scale-up and 

integration of research areas that have continued to make individual progress. Significant 

improvements have been made in the key areas of reforming and alcohol synthesis. Pilot-scale 

equipment to demonstrate continuous fluidized bed reforming and alcohol synthesis is expected 

to soon be operational at NREL. 

Once test facilities become available, a significant effort will be required to demonstrate 

improved longevity of tar reforming and alcohol synthesis catalysts. While catalyst improvement 

is one path for achieving these goals, the identification of optimal operating conditions for 

existing catalysts will also play a significant role. 

Research and development activities to improve the performance of both mixed alcohol and tar 

reforming catalysts are ongoing and are leveraged through collaboration with several industrial 

partners (names withheld for confidentiality). Improvements will be captured in the process 

model, and translated to full analysis of MESP, in future State of Technology (SOT) reports. 

59
 
 

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Renewable Energy Laboratory, 2009. http://www.nrel.gov/docs/fy09osti/45913.pdf. 

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Hydrogen,” Chapter 9.2. Rosillo-Calle, F., et al., ed. Industrial Uses of Biomass Energy. 

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11. Rodrigues, A.P. Participação dos fornecedores de cana na cadeia do açúcar e álcool. 

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12. Seabra, J.E.A. Technical-Economic Evaluation of Options for Whole Use of Sugar Cane 

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Estadual de Campinas, 2007 (in Portuguese). 

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13. Graboski, M.S.; McCormick, R.L. “Combustion of Fat and Vegetable Oil Derived Fuels in 

Diesel Engines.” Prog. Energy Combust. Sci.; Vol. 24, 1998; pp. 125-164. 

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Production Cost.” Biores. Technol.; Vol. 97, 2006; pp. 671-678. 

15. Garrett, D.E., Chemical Engineering Economics. New York: Van Nostrand Reinhold, 1989. 

16. Peters, M.S.; Timmerhaus, K.D.; West, R. Plant Design and Economics for Chemical 

Engineers, 5th ed., International Edition. New York: McGraw-Hill, Inc., 2004. 

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18. Plant Cost Index. Chemical Engineering. December 2010. 

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a Billion-Ton Annual Supply. DOE/GO-102995-2135 & ORNL/TM-2005/66. A joint U.S. 

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20. Craig, K.R.; Mann, M.K. Cost and Performance Analysis of Biomass-Based Integrated 

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Residues to a Methane-Rich Gas in a High-Throughput Gasifier. Report DE88013138. 

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

L.; Slayton, A.; Lukas, J. Lignocellulosic Biomass to Ethanol Process Design and Economics 

Utilizing Co-Current Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for Corn Stover. 

Report No. NREL/TP-510-32438. Golden, CO: National Renewable Energy Laboratory, 

June 2002. http://www.nrel.gov/docs/fy02osti/32438.pdf. 

23. Magee, J.S. Fluid Catalytic Cracking: Science and Technology. The Netherlands: Elsevier 

Publishing Company, July 1993; p. 21. 

24. Apanel, G.; Wright, H.A. “System and Method for Dual Fluidized Bed Gasification,” U.S. 

Patent Application No. US 2010/0181539 A1, July 22, 2010. 

25. Cohen, A.P.; Reynolds, T.M.; Davis, M.M. “Adsorbent for Drying Ethanol,” U.S. Patent 

Application No. US 2010/0081851 A1, April 1, 2010. 

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26. Pottie, M.; Guimier, D. Preparation of Forest Biomass for Optimal Conversion. PointeClaire, Quebec, Canada: FERIC IEA (Forest Engineering Research Institute of Canada 

International Energy Agency), 1985. ISSN 0381-7733. 

27. Rummer, Dr. R., Project Leader, Forest Operations Research, USDA Forest Service, Auburn, 

AL, personal communication with the INL authors. 

28. Hall, P. Storage Guidelines for Wood Residues for Bioenergy. Prepared by Scion Next 

Energy Biomaterials for EECA (Energy Efficiency and Conservation Authority), 2009. 

29. Sinclair, S.; Hassler, C.; Bolstad, K. “Moisture Loss in Aspen Logging Residue.” Wood and 

Fiber Science; Vol. 16, No. 1, 1984; pp. 93-96. 

30. Mann, M.K.; Spath, P.L. Life Cycle Assessment of a Biomass Gasification Combined-Cycle 

Power System. Report No. NREL/TP-430-23076. Golden, CO: National Renewable Energy 

Laboratory, 1997. http://www.nrel.gov/docs/legosti/fy98/23076.pdf. 

31. Paisley, M.A. “Advanced Biomass Gasification for the Production of Biopower, Fuels and 

Chemicals.” Presented at the 2007 AIChE Annual Meeting, Salt Lake City, Utah. 

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Pomeroy, M. D.; Phillips, S. D. “Demonstration and Characterization of Ni/Mg/K/AD90 

Used for Pilot-Scale Conditioning of Biomass-Derived Syngas.” Catalysis Letters; Vol. 

134(3-4), 2010; pp. 242-249. 

33. Ravi, G.; Gupta, S.K.; Viswanathan, S.; Ray, M.B. “Optimization of Venturi Scrubbers 

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34. Doctor, R.D.; Molburg, J.C.; Thimmapuram, P.R. KRW Oxygen-Blown Gasification 

Combined Cycle: Carbon Dioxide Recovery, Transport, and Disposal. Argonne, Illinois: 

Argonne National Laboratory Energy Systems Division, August 1996. http://www.fischertropsch.org/DOE/DOE_reports/Doctor/ANL-ESD-34/ANL-ESD34,%20Part%201,%20Pages%201-83.pdf. 

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York: Marcel Dekker, Inc., 2001. 

36. “Sylobead Desiccant Process Adsorbents for Ethanol Enrichment.” Grace Davison BioFuel 

Technologies Group Products website. 

http://www.grace.com/Products/Sylobead/Default.aspx. 

37. UOP Ethanol Dehydration Solutions. UOP 4979. Des Plaines, IL: UOP LLC, June 2008. 

http://www.uop.com/wp-content/uploads/2011/02/UOP-Ethanol-Dehydration-SolutionsBrochure.pdf. 

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38. Humbird, D.; Davis, R.; Tao, L.; Kinchin, C.; Hsu, D.; Aden, A.; Schoen, P.; Lukas, J.; 

Olthof, B.; Worley, M.; Sexton, D.; Dudgeon, D. Process Design and Economics for 

Biochemical Conversion of Lignocellulosic Biomass to Ethanol: Dilute-Acid Pretreatment 

and Enzymatic Hydrolysis of Corn Stover. Report No. TP-5100-47764. Golden, CO: National 

Renewable Energy Laboratory, 2011. http://www.nrel.gov/docs/fy11osti/47764.pdf. 

39. GPSA. Engineering Data Book, FPS Version, 12th ed. Tulsa, OK: Gas Processors Suppliers 

Association, 2004. 

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Press, 1981. 

41. Keeney, D.; Muller, M. Water Use by Ethanol Plants: Potential Challenges. Minneapolis, 

MN: Institute for Agriculture and Trade Policy, October 2006. 

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Ethanol and Petroleum Gasoline. Report No. ANL/ESD/09-1. Argonne, IL: Center for 

Transportation Research, Energy Systems Division, Argonne National Laboratory, January 

2009. http://www.transportation.anl.gov/pdfs/AF/557.pdf. 

43. Jeakel, D., price quote from AGSCO for super sacks or bulk, 2004. 

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45. “Annual Energy Outlook 2011.” Data Table A3 – Energy Prices by Sector and Source. 

Energy Information Administration (EIA) website, 

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Cost. Pittsburgh, PA: Carnegie Mellon University, 2004. 

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Products.” Chemical Economics Handbook: The Economic Environment of the Chemical 

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Prices. Energy Information Administration (EIA) website, March 2009, 

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63
 
64 

Appendix A. Equipment Description and Design Basis Summary 
 

 

 

NREL’s Process Engineering Team has compiled information on the equipment in the 

benchmark model. The compilation contains information about the cost, reference year, scaling 

factor, scaling characteristic, design information, and back‐up cost referencing. The information 

is stored in the economic spreadsheet portion of the model. 

The following table summarizes the important fields of information contained in the spreadsheet 

model. A partial listing of the information is attached for each piece of equipment. Additional 

information from the spreadsheet model is presented in Appendix B. 

Equipment number:a,b	 Unique identifier, the first letter indicates the equipment type and the first number 

represents the process area, e.g., P-601 is a pump in Area 600 

Equipment name:a,b	 Descriptive name of the piece of equipment 

Associated PFD:	 PFD number on which the piece of equipment appears, e.g., PFD-P820-A101 

Equipment category:a	 Code indicating the general type of equipment, e.g., PUMP 

Equipment type:a	 Code indicating the specific type of equipment, e.g., CENTRIFUGAL for a pump 

Equipment description:a	 Short description of the size or characteristics of the piece of equipment, e.g., 20 

gpm, 82 ft head for a pump 

Number required:b	 Number of duplicate pieces of equipment needed 

Number spares:b	 Number of on-line spares 

Scaling stream:b	 Stream number or other characteristic variable from the Aspen Plus model by 

which the equipment cost is be scaled 

Base cost:b	 Equipment cost 

Cost basis:a	 Source of the equipment cost, e.g., Aspen Icarus 2006.5 or vendor quotation 

Cost year:b	 Year in which the cost estimate is based 

Base for scaling:b	 Value of the scaling stream or variable used to obtain the base cost of the 

equipment 

Base type:	 Type of variable used for scaling, e.g., FLOW, DUTY, etc. 

Base units:	 Units of the scaling stream or variable, e.g., LB/HR, MMBTU/HR 

Installation factor:b	 Value of the installation factor. Installed Cost = Base Cost x Installation Factor 

Installation factor basis:	 Source of the installation factor, e.g., Aspen Icarus 2006.5, vendor quotation 

Scale factor exponent:b	 Value of the exponential scaling equation 

Scale factor basis:	 Source of the scaling exponent value, e.g., vendor quotation 

Material of construction:a	 Material of construction 

Notes:	 Any other important information about the design or cost 

Document:	 Complete, multi-page document containing design calculations, vendor literature 

and quotations, and any other important information. This is stored as an electronic 

document and can be pages from a spreadsheet, other electronic sources, or 

scanned information from vendors. 

Design date:	 Original date for the design of this piece of equipment 

Modified date:	 The system automatically marks the date in this field whenever any field is 

changed 

a These fields are listed for all pieces of equipment in Appendix A. 

b These fields are part of the equipment cost listing in Appendix B. 

65
 
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A101 to A102

C‐101 Hopper Feeder CONVEYOR VIBRATING‐FEEDER Included in per unit cost of feestock per Idaho National Laboratory (INL) 

C‐102 Screener Feeder Conveyor CONVEYOR BELT Included in per unit cost of feestock per Idaho National Laboratory (INL) 

C‐103 Radial Stacker Conveyor CONVEYOR BELT Included in per unit cost of feestock per Idaho National Laboratory (INL) 

C‐104 Dryer Feed Screw Conveyor CONVEYOR SCREW Included in per unit cost of feestock per Idaho National Laboratory (INL) 

C‐105 Gasifier Feed Screw Conveyor CONVEYOR SCREW Included in per unit cost of feestock per Idaho National Laboratory (INL) 

K‐101 Flue Gas Blower FAN CENTRIFUGAL Included in per unit cost of feestock per Idaho National Laboratory (INL) 

K‐102 CO2 Booster Blower COMPRESSOR CENTRIFUGAL Included in per unit cost of feestock per Idaho National Laboratory (INL) M‐101 Hydraulic Truck Dump with Scale SCALE TRUCK‐SCALE Included in per unit cost of feestock per Idaho National Laboratory (INL) 

M‐103 Front End Loaders VEHICLE LOADER Included in per unit cost of feestock per Idaho National Laboratory (INL) 

M‐104 Waste Heat Dryer DRYER ROTARY‐DRUM Included in per unit cost of feestock per Idaho National Laboratory (INL) 

S‐101 Magnetic Head Pulley SEPARATOR MAGNET Included in per unit cost of feestock per Idaho National Laboratory (INL) 

S‐102 Vibratory Conveyor SEPARATOR SCREEN Included in per unit cost of feestock per Idaho National Laboratory (INL) 

T‐101 Dump Hopper TANK LIVE‐BTM‐BIN Included in per unit cost of feestock per Idaho National Laboratory (INL) 

T‐103 Dryer Feed Bin TANK LIVE‐BTM‐BIN Included in per unit cost of feestock per Idaho National Laboratory (INL) 

T‐104 Dried Biomass Hopper TANK VERTICAL‐VESSEL Included in per unit cost of feestock per Idaho National Laboratory (INL) 

T‐105 Lock Hopper TANK VERTICAL‐VESSEL Included in per unit cost of feestock per Idaho National Laboratory (INL) 

T‐106 Feed Hopper TANK VERTICAL‐VESSEL Included in per unit cost of feestock per Idaho National Laboratory (INL) 

66
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A201 to A202

C‐201 Sand / Ash Conditioner / Conveyor CONVEYOR SCREW Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

H‐201 Syngas Quench Inlet / Combustion Air Pre‐Heat Exchanger * HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 4.64; LMTD (°F) = 14.48; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 12829; Included in Tay

H‐202 Tar Reformer Effluent / Combustion Air Pre‐Heat Exchanger * HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 9.95; LMTD (°F) = 24.44; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 20350; Included in Tay

K‐202 Char Combustor Air Blower COMPRESSOR CENTRIFUGAL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

M‐201 Sand / Ash Cooler MISCELLANEOUS MISCELLANEOUS Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

R‐201 Indirectly‐Heated Biomass Gasifier REACTOR VERTICAL‐VESSEL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

R‐202 Char Combustor REACTOR VERTICAL‐VESSEL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

S‐201 Primary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

S‐202 Secondary Gasifier Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

S‐203 Primary Combustor Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

S‐204 Char Combustor Secondary Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

T‐201 Sand / Ash Bin TANK FLAT‐BTM‐STORAGE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

67
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A301 to A306

H‐301 Quench Water Recirculation Cooler HEATX SHELL‐TUBE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass EnergyH‐302 Steam Blowdown / Unreacted Syngas Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 4.38; LMTD (°F) = 43; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 3395 

H‐303 Syngas Quench Inlet / Make‐Up Water Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 7.22; LMTD (°F) = 15.11; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 15919 H‐304 Syngas Quench Inlet / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 5.66; LMTD (°F) = 92.05; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 2461 

H‐305 Alcohol Synthesis Effluent / Combustion Air Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 1.58; LMTD (°F) = 44.04; U (Btu / Hr‐ft^2‐°F)=40; A (ft^2) = 894 

H‐306 Syngas Quench Inlet / Combustion Air Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 2.64; LMTD (°F) = 44.04; U (Btu / Hr‐ft^2‐°F)=40; A (ft^2) = 1497 

H‐307 Syngas Quench Inlet / Combustion Air Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 5.91; LMTD (°F) = 43.18; U (Btu / Hr‐ft^2‐°F)=40; A (ft^2) = 3422 

H‐308 Tar Reformer Effluent / Combustion Air Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 5.33; LMTD (°F) = 24.44; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 10910 

H‐309 Tar Reformer Effluent / Combustion Air Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 12.75; LMTD (°F) = 322.72; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 1975 

H‐310 Catalyst Regenerator Flue Gas / Tar Reformer Steam & Recycle Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 69.86; LMTD (°F) = 244.39; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 14293 

H‐311 Tar Reformer Effluent / Unreacted Syngas Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 9.08; LMTD (°F) = 397.35; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 1142 

K‐305 Combustion Air Blower COMPRESSOR CENTRIFUGAL Volumetric Inlet Flow Rate = 66,000 CFM; Design Inlet Pressure = 0 PSIG; Design Outlet Pressure = 5 PS

M‐300 Combustor Fuel Pre‐Mix MISCELLANEOUS MISCELLANEOUS Piping included in installation factor M‐300H Tar Reformer Inlet Mix MISCELLANEOUS MISCELLANEOUS Piping included in installation factor 

M‐301 Syngas Cyclone Separator MISCELLANEOUS VERTICAL‐VESSEL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

M‐302 Syngas Venturi Scrubber MISCELLANEOUS VERTICAL‐VESSEL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass EnergyP‐302 Quench Water Recirculation Pump PUMP CENTRIFUGAL Flow = 1,010 GPM; Fluid Head = 60 Ft; Design Pressure = 30 PSIG 

R‐301A Tar Reformer Catalyst Regenerator REACTOR VERTICAL‐VESSEL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

R‐303 Tar Reformer REACTOR VERTICAL‐VESSEL Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

S‐306 Tar Reformer 2‐Stage Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

S‐307 Catalyst Regenerator 2‐Stage Cyclone SEPARATOR GAS CYCLONE Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

68
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A401 to A408

H‐401 Alcohol Synthesis Feed / Effluent Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 127.7; LMTD (°F) = 30.82; U (Btu / Hr‐ft^2‐°F)=90; A (ft^2) = 46034 

H‐402 Alcohol Synthesis Feed / Effluent Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 73.55; LMTD (°F) = 16.57; U (Btu / Hr‐ft^2‐°F)=90; A (ft^2) = 49332 

H‐403 Compressor Interstage / Deaerator Feed Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 25.04; LMTD (°F) = 25.9; U (Btu / Hr‐ft^2‐°F)=35; A (ft^2) = 27629 

H‐404 Compressor Interstage / Deaerator Feed Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 15.3; LMTD (°F) = 13.88; U (Btu / Hr‐ft^2‐°F)=50; A (ft^2) = 22037 

H‐405 Compressor Interstage / Deaerator Feed Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 15.53; LMTD (°F) = 13.4; U (Btu / Hr‐ft^2‐°F)=100; A (ft^2) = 11595 

H‐406 Compressor Interstage / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 15.96; LMTD (°F) = 42.34; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 12567 

H‐407 Compressor Interstage / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 11.1; LMTD (°F) = 42.34; U (Btu / Hr‐ft^2‐°F)=50; A (ft^2) = 5244 

H‐408 Compressor Interstage / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 11.39; LMTD (°F) = 42.34; U (Btu / Hr‐ft^2‐°F)=70; A (ft^2) = 3844 

H‐409 Alcohol Synthesis Effluent / LO‐CAT Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 0.05; LMTD (°F) = 70.58; U (Btu / Hr‐ft^2‐°F)=50; A (ft^2) = 14 

H‐410A 1st Stage Air Intercooler HEATX AIR‐COOLED EXCHANGER Duty = 37.4 MMBtu/Hr; Bare Tube Surface Area = 11,509 Ft^2; Design Pressure = 40 PSIG; OverdesigH‐410B 2nd Stage Air Intercooler HEATX AIR‐COOLED EXCHANGER Duty = 36.0 MMBtu/Hr; Bare Tube Surface Area = 8,613 Ft^2; Design Pressure = 150 PSIG; Overdesig

H‐410C 3rd Stage Air Intercooler HEATX AIR‐COOLED EXCHANGER Duty = 32.6 MMBtu/Hr; Bare Tube Surface Area = 6,891 Ft^2; Design Pressure = 500 PSIG; Overdesig

H‐410D 4th Stage Air Intercooler HEATX AIR‐COOLED EXCHANGER Duty = 38.6 MMBtu/Hr; Bare Tube Surface Area = 6,891 Ft^2; Design Pressure = 1,500 PSIG; Overdes

H‐410E 5th Stage Air Intercooler HEATX AIR‐COOLED EXCHANGER Duty = 36.0 MMBtu/Hr; Bare Tube Surface Area = 5,507 Ft^2; Design Pressure = 3,300 PSIG; Overdes

H‐411A 1st Stage Water Intercooler HEATX SHELL‐TUBE Duty = 11.1 MMBtu/Hr; LMTD = 20.0 F; U = 40 Btu/hr‐ft^2‐F; Surface Area = 13,875 Ft^2; Fixed TS; D

H‐411B 2nd Stage Water Intercooler HEATX SHELL‐TUBE Duty = 6.4 MMBtu/Hr; LMTD = 20.0 F; U = 55 Btu/hr‐ft^2‐F; Surface Area = 5,818 Ft^2; Fixed TS; De

H‐411C 3rd Stage Water Intercooler HEATX SHELL‐TUBE Duty = 4.4 MMBtu/Hr; LMTD = 20.0 F; U = 70 Btu/hr‐ft^2‐F; Surface Area = 3,143 Ft^2; Fixed TS; De

H‐411D 4th Stage Water Intercooler HEATX SHELL‐TUBE Duty = 3.3 MMBtu/Hr; LMTD = 20.0 F; U = 85 Btu/hr‐ft^2‐F; Surface Area = 1,941 Ft^2; Fixed TS; De

H‐411E 5th Stage Water Intercooler HEATX SHELL‐TUBE Duty = 3.3 MMBtu/Hr; LMTD = 20.0 F; U = 95 Btu/hr‐ft^2‐F; Surface Area = 1,650 Ft^2; Fixed TS; De

H‐413 Syngas Air Cooler HEATX AIR‐COOLED EXCHANGER Duty = 33.8 MMBtu/Hr; Bare Tube Surface Area = 11,438 Ft^2; Design Pressure = 3,300 PSIG 

H‐414 Syngas Cooling Water Exchanger HEATX SHELL‐TUBE Duty = 15.6 MMBtu/Hr; LMTD = 24.7 F; U = 95 Btu/hr‐ft^2‐F; Surface Area = 6,658 Ft^2; Fixed TS; D

H‐415 Alcohol Synthesis Effluent / Syngas Recycle Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 19.56; LMTD (°F) = 36.34; U (Btu / Hr‐ft^2‐°F)=90; A (ft^2) = 5980 

H‐416 Alcohol Synthesis Effluent / CO2‐Rich Gas Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 1.01; LMTD (°F) = 58.47; U (Btu / Hr‐ft^2‐°F)=80; A (ft^2) = 215 

H‐417 Alcohol Synthesis Effluent / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 5.35; LMTD (°F) = 59.33; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 3608 

H‐418 Alcohol Synthesis Effluent / Expander Inlet Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 9.71; LMTD (°F) = 62.15; U (Btu / Hr‐ft^2‐°F)=80; A (ft^2) = 1953 

H‐419 Tar Reformer Effluent / Expander Inlet Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 3.81; LMTD (°F) = 24.44; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 5202 

H‐420 Tar Reformer Effluent / Expander Inlet Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 25.07; LMTD (°F) = 583.69; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 1432 

H‐421 Alcohol Synthesis Effluent / Expander Inlet Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 4.31; LMTD (°F) = 78.57; U (Btu / Hr‐ft^2‐°F)=70; A (ft^2) = 783 

H‐422 Tar Reformer Effluent / Expander Inlet Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 10.97; LMTD (°F) = 630.9; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 695 

H‐423 Tar Reformer Effluent / Expander Inlet Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 16.3; LMTD (°F) = 459.7; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 1418 

H‐424 Alcohol Synthesis Effluent / CO2‐Rich Gas Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 2.37; LMTD (°F) = 44.27; U (Btu / Hr‐ft^2‐°F)=45; A (ft^2) = 1192 

H‐425 Char Combustor Flue Gas / CO2‐Rich Gas Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 4.52; LMTD (°F) = 550.38; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 410 

H‐426 Tar Reformer Effluent / CO2‐Rich Gas Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 10.14; LMTD (°F) = 524.35; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 967 

H‐427 Catalyst Regenerator Flue Gas / Alcohol Synthesis Feed Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 20.42; LMTD (°F) = 364.97; U (Btu / Hr‐ft^2‐°F)=40; A (ft^2) = 1399 

K‐410 Synthesis Gas Compressor COMPRESSOR CENTRIFUGAL Quotation from Elliott inlcudles multi‐stage centrifugal compressor system, motors and variable frequ

K‐414 Unconverted Syngas Recycle Compressor COMPRESSOR CENTRIFUGAL Cost for recycle compressor service included in quotation from Elliott 

K‐412A Purge Gas Expander Stage 1 COMPRESSOR CENTRIFUGAL Inlet Flow Rate = 795 CFM; Power Recovered = 10,058 HP 

K‐412B Purge Gas Expander Stage 2 COMPRESSOR CENTRIFUGAL Inlet Flow Rate = 4,959 CFM; Power Recovered = 12,257 HP 

M‐401 DEPG Acid Gas Removal System MISCELLANEOUS PACKAGE UNIT Detailed cost estimate for system is provided in Appendices M‐402 Ammonia Absorption Refrigeration System MISCELLANEOUS PACKAGE UNIT Equipment sizing and cost estimate based on refrigeration demands of DEPG and Methanol / H2S syst

69
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

M‐403 Amine Acid Gas Enrichment System MISCELLANEOUS PACKAGE UNIT Detailed cost estimate for system is provided in Appendices 

M‐404 Methanol / H2S Absorption System MISCELLANEOUS PACKAGE UNIT Detailed cost estimate for system is provided in Appendices 

M‐405 LO‐CAT Sulfur Recovery System MISCELLANEOUS PACKAGE UNIT Estimated cost was interpolated from other projects based on sulfur recovery rate and H2S concentrat

R‐410 Alcohol Synthesis Reactor REACTOR TUBULAR REACTOR Tubular Reactor. Design Pressure = 3,300 PSIG; Tubes / Shell = 4500; Tube ID = 1.5 in; Tube Length = S‐410A 1st Interstage KO Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 225% of Critical Velocity; Diameter = 12 Ft; T‐T Height = 24 Ft; Design Pressure = 40 P

S‐410B 2nd Interstage KO Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 225% of Critical Velocity; Diameter = 9 Ft; T‐T Height = 18 Ft; Design Pressure = 150 P

S‐410C 3rd Interstage KO Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 225% of Critical Velocity; Diameter = 6 Ft; T‐T Height = 15 Ft; Design Pressure = 500 P

S‐410D 4th Interstage KO Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 225% of Critical Velocity; Diameter = 5 Ft; T‐T Height = 12 Ft; Design Pressure = 1,500 

S‐410E 5th Interstage KO Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 225% of Critical Velocity; Diameter = 4 Ft; T‐T Height = 12 Ft; Design Pressure = 3,300 

S‐415 Pre‐Compressor KO Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 225% of Critical Velocity; Diameter = 16 Ft; T‐T Height = 32 Ft; Design Pressure = 10 P

70
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A501 to A503

D‐504 Crude Alcohol Distillation Column COLUMN DISTILLATION 60 Theoretical Stages at 60% Efficiency = 100 Trays; 24 in Tray Spacing; Diameter = 9 ft; T‐T Height 

D‐505 Methanol Column COLUMN DISTILLATION 81 Theoretical Stages at 60% Efficiency = 135 Trays; 24 in Tray Spacing; Diameter = 12 ft; T‐T Height = 

H‐501 Ethanol Product / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 4.81; LMTD (°F) = 30.18; U (Btu / Hr‐ft^2‐°F)=90; A (ft^2) = 1770 

H‐502 Higher Alcohols Product / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 0.76; LMTD (°F) = 51.77; U (Btu / Hr‐ft^2‐°F)=90; A (ft^2) = 164 

H‐503 Syngas Quench Inlet / Water‐Methanol Recycle Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 12.14; LMTD (°F) = 17.64; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 22949 

H‐504C Crude Alcohol Column Condenser HEATX AIR‐COOLED EXCHANGER Duty = 36.0 MMBtu/Hr; Bare Tube Surface Area = 8,406 Ft^2; Design Pressure = 50 PSIG 

H‐504R Crude Alcohol Column Reboiler HEATX SHELL‐TUBE Duty = 33.4 MMBtu/Hr; LMTD = 30 F; U = 150 Btu/hr‐ft^2‐F; Surface Area = 7,480 Ft^2; U‐Tube Ket

H‐505C Methanol Column Condenser HEATX AIR‐COOLED EXCHANGER Duty = 100.3 MMBtu/Hr; Bare Tube Surface Area = 23,460 Ft^2; Design Pressure = 50 PSIG 

H‐505R Methanol Column Reboiler HEATX SHELL‐TUBE Duty = 84.1 MMBtu/Hr; LMTD = 30 F; U = 150 Btu/hr‐ft^2‐F; Surface Area = 18,860 Ft^2; U‐Tube Ke

H‐506 Syngas Quench Inlet / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 10.42; LMTD (°F) = 47.31; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 8809 H‐507 Alcohol Synthesis Effluent / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 1.03; LMTD (°F) = 40.7; U (Btu / Hr‐ft^2‐°F)=50; A (ft^2) = 505 

H‐508 Alcohol Synthesis Effluent / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 0.88; LMTD (°F) = 64.96; U (Btu / Hr‐ft^2‐°F)=40; A (ft^2) = 340 

H‐509 Catalyst Regenerator Flue Gas / Water‐Methanol Recycle Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 0.48; LMTD (°F) = 509.13; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 47 

H‐510 Tar Reformer Effluent / Water‐Methanol Recycle Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 1.52; LMTD (°F) = 364.71; U (Btu / Hr‐ft^2‐°F)=20; A (ft^2) = 208 

H‐513 Molecular Sieve Flush Cooler HEATX AIR‐COOLED EXCHANGER Duty = 36.0 MMBtu/Hr; Bare Tube Surface Area = 8,406 Ft^2; Design Pressure = 50 PSIG 

H‐514 Methanol Condenser HEATX AIR‐COOLED EXCHANGER Duty = 36.0 MMBtu/Hr; Bare Tube Surface Area = 8,406 Ft^2; Design Pressure = 50 PSIG 

H‐591 Higher Alcohols Product Finishing Cooler HEATX SHELL‐TUBE Duty = 0.1 MMBtu/Hr; LMTD = 15 F; U = 100 Btu/hr‐ft^2‐F; Surface Area = 30 Ft^2; U‐Tube; Design 

H‐593 Ethanol Product Finishing Cooler HEATX SHELL‐TUBE Duty = 0.4 MMBtu/Hr; LMTD = 15 F; U = 100 Btu/hr‐ft^2‐F; Surface Area = 265 Ft^2; U‐Tube; Desig

P‐504 Crude Alcohol Column Reflux Pump PUMP CENTRIFUGAL Flow Rate = 200 gpm; Liquid Head = 50 ft 

P‐505 Methanol Column Reflux Pump PUMP CENTRIFUGAL Flow Rate = 550 gpm; Liquid Head = 50 ft 

P‐514 Condensed Methanol Pump PUMP CENTRIFUGAL Flow = 120 GPM; Fluid Head = 9,200 Ft; Design Pressure = 4,000 PSIG 

P‐590 Crude Alcohol Column Bottoms Pump PUMP CENTRIFUGAL Flow Rate = 300 gpm; Liquid Head = 70 ft 

P‐592 Methanol Column Bottoms Pump PUMP CENTRIFUGAL Flow Rate = 900 gpm; Liquid Head = 70 ft 

S‐501 Mixed Alcohols Condensation Knock‐out SEPARATOR KNOCK‐OUT DRUM Horizontal Drum; Diameter = 5 ft; T‐T Length = 10 ft; Design Pressure = 3,300 PSIG 

S‐502 Mol Sieve Pre‐Flash Drum SEPARATOR KNOCK‐OUT DRUM Vertical Drum; 100% of Critical Velocity; Diameter = 10 ft; T‐T Length = 30 ft; Design Pressure = 35 PSS‐503 Molecular Sieve Separator System MISCELLANEOUS PACKAGE UNIT Superheater, twin mole sieve columns, product cooler, condenser, pumps, vacuum source. 

S‐514 Condensed Methanol Flash Drum SEPARATOR VERTICAL‐VESSEL Vertical Drum; 100% of Critical Velocity; Diameter = 2 Ft; T‐T Height = 10 Ft; Design Pressure = 50 PS

T‐504 Crude Alcohol Column Overhead Accumulator SEPARATOR HORIZONTAL‐VESSEL Reflux Flow = 200 gpm; 10 Minute Liquid Inventory; 2,000 Gallon Liquid Inventory; Vapor Space Fact

T‐505 Methanol Column Overhead Accumulator SEPARATOR HORIZONTAL‐VESSEL Reflux Flow = 550 gpm; 10 Minute Liquid Inventory; 5,500 Gallon Liquid Inventory; Vapor Space Fact

71
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A601 to A604

H‐601 Steam Turbine Condenser HEATX AIR‐COOLED EXCHANGER Modular units with cost of $1.3 MM per unit. Cooling capacity of 58.25 MMBtu / Hr per module. 

H‐602 Steam Turbine Exhaust / Syngas Recycle Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 3.92; LMTD (°F) = 49.3; U (Btu / Hr‐ft^2‐°F)=40; A (ft^2) = 1986 

H‐603 Blowdown Cooler HEATX SHELL‐TUBE Duty = 0.4 MMBtu/Hr; LMTD = 15 F; U = 100 Btu/hr‐ft^2‐F; Surface Area = 265 Ft^2; U‐Tube; Desig

H‐604 Steam Turbine Exhaust / Mol‐Sieve Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 4.38; LMTD (°F) = 70.77; U (Btu / Hr‐ft^2‐°F)=35; A (ft^2) = 1767 

H‐605 Syngas Quench Inlet / Make‐Up Water Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 6.5; LMTD (°F) = 138.95; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 1559 

H‐606 Syngas Quench Inlet / BFW Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 6.5; LMTD (°F) = 23.48; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 9232 

H‐607 Alcohol Synthesis Effluent / BFW Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 9.98; LMTD (°F) = 23.48; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 14178 

H‐608 Tar Reformer Effluent / BFW Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 63.73; LMTD (°F) = 139.77; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 15198 

H‐609 Catalyst Regenerator Flue Gas / BFW Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 26.06; LMTD (°F) = 203.99; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 4259 

H‐610 Char Combustor Flue Gas / BFW Pre‐Heat Exchanger HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 38.11; LMTD (°F) = 195.12; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 6511 

H‐611 Catalyst Regenerator Flue Gas Steam Generator HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 17.67; LMTD (°F) = 353.55; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 1666 

H‐612 Char Combustor Flue Gas Steam Superheater HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 147.91; LMTD (°F) = 441.64; U (Btu / Hr‐ft^2‐°F)=30; A (ft^2) = 11163 

H‐613 Tar Reformer Effluent Steam Superheater HEATX SHELL‐TUBE Duty (MMBtu / Hr) = 74.62; LMTD (°F) = 362.62; U (Btu / Hr‐ft^2‐°F)=25; A (ft^2) = 8231 

M‐601A BFW Reverse Osmosis (RO) Unit MISCELLANEOUS PACKAGE UNIT 600 gpm package unit. Includes RO units, electrodeionization system, CIP skid, brine recovery and con

M‐601B BFW Electrodeionization (EDI) Unit MISCELLANEOUS PACKAGE UNIT 600 gpm package unit. Includes RO units, electrodeionization system, CIP skid, brine recovery and con

M‐601C Hot Condensate Polishing Unit MISCELLANEOUS PACKAGE UNIT 600 gpm package unit. Includes RO units, electrodeionization system, CIP skid, brine recovery and con

M‐602A Extraction Steam Turbine / Generator Stage 1 GENERATOR STEAM‐TURBINE Estimate based on vendor quotation for 30.14 MW system. 

M‐602B Extraction Steam Turbine / Generator Stage 2 GENERATOR STEAM‐TURBINE Estimate based on vendor quotation for 30.14 MW system. Includes asociated equipment with except

M‐603 Startup Boiler MISCELLANEOUS PACKAGE UNIT Assume steam requirement equal to 1/2 of steam requirement for gasifier at full rate steam rate = 36,

M‐604 Brine Recovery Reverse Osmosis (RO) Unit MISCELLANEOUS PACKAGE UNIT 600 gpm package unit. Includes RO units, electrodeionization system, CIP skid, brine recovery and con

P‐601 Make‐Up Pump PUMP CENTRIFUGAL Flow Rate = 160 gpm; Liquid Head = 35 ft 

P‐602 Condensate Pump PUMP CENTRIFUGAL Flow Rate = 500 gpm; Liquid Head = 60 ft 

P‐603 EDI Pump PUMP CENTRIFUGAL Flow Rate = 500 gpm; Liquid Head = 60 ft 

P‐604 Boiler Feed Water Pump PUMP POSITIVE DISPLACEMENT Flow Rate = 1,050 gpm; Liquid Head = 3,340 ft 

S‐601 Blowdown Flash Drum TANK VERTICAL‐VESSEL Liquid Flow Rate = 30 gpm; 1,500 psig Design Pressure; 10 Minute Residence Time; Vapor Space Fact

T‐601 Condensate Collection Tank SEPARATOR VERTICAL‐VESSEL Condensate Flow = 1,000 gpm; Residence Time = 5 Minutes; Liquid Volume = 5,000 Gallons T‐602 Condensate Surge Tank TANK HORIZONTAL‐VESSEL Condensate Flow = 1,000 gpm; Residence Time = 5 Minutes; Liquid Volume = 5,000 Gallons 

T‐603 Deaerator TANK HORIZONTAL‐VESSEL Liquid Flow Rate = 1,050 gpm; 50 psig Design Pressure; 10 Minute Residence Time; Vapor Space Fact

T‐603A Deaerator Packed Column COLUMN PACKED COLUMN Inside Diameter = 2 ft; Height = 10 ft; Packing Height = 8 ft 

T‐605 Steam Drum TANK HORIZONTAL‐VESSEL Horizontal Drum; < 100% of Critical Velocity; Diameter = 5 ft; T‐T Length = 20 ft 

72
EQUIPMENT_NUM EQUIPMENT_NAME EQUIPMENT_CATEGORY EQUIPMENT_TYPE EQUIPMENT_DESCRIPTION 

PFD‐P820‐A701 to A702

K‐701 Plant Air Compressor COMPRESSOR RECIPROCATING Flow Rate = 450 CFM; Outlet Pressure = 125 PSIG 

M‐701 Cooling Tower System COOLING‐TOWER INDUCED‐DRAFT 15,000 gpm, 110 °F Cooling Tower Inlet Temperature 

M‐702 Hydraulic Truck Dump with Scale SCALE TRUCK‐SCALE Hydraulic Truck Dumper with Scale 

M‐703 Flue Gas Scrubber MISCELLANEOUS VERTICAL‐VESSEL Scrubber costs from TurboSonic Technologies, Inc. Quotation basis = 160,000 ACF / Minute 

M‐704 Flue Gas Stack MISCELLANEOUS MISCELLANEOUS Base Diameter = 72 inches; Stack Height = 200 ft 

M‐705 Clarifier SEPARATOR CLARIFIER Sizing Basis = 1.4 Ft / Hr Rise Rate; 40' Diameter x 12' Height 

M‐706 Belt Press MISCELLANEOUS MISCELLANEOUS Estimate from Harris Group Inc. 

M‐707 Sand Filter MISCELLANEOUS MISCELLANEOUS Estimate from Harris Group Inc. 

M‐708 Emergency Flare MISCELLANEOUS MISCELLANEOUS Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

M‐709 Cooling Tower Reverse Osmosis (RO) Unit MISCELLANEOUS PACKAGE UNIT 600 gpm package unit. Includes RO units, electrodeionization system, CIP skid, brine recovery and con

M‐710 Product Loading Rack MISCELLANEOUS MISCELLANEOUS Assumed to be included in Site Development and Equipment Installation 

P‐701 Cooling Water Pump PUMP CENTRIFUGAL Flow Rate = 18,620 gpm; Liquid Head = 185 ft 

P‐702 Firewater Pump PUMP CENTRIFUGAL Flow Rate = 2,500 gpm; Liquid Head = 120 ft 

P‐703 Diesel Pump PUMP CENTRIFUGAL Flow Rate = 5 gpm; Liquid Head = 150 ft 

P‐704 Ammonia Pump PUMP CENTRIFUGAL Flow Rate = 1 gpm; Liquid Head = 150 ft 

P‐705 Caustic Pump PUMP CENTRIFUGAL Flow = 0.5 GPM; Fluid Head = 10 Ft; Design Pressure = 150 PSIG 

P‐707 BFW Chemical Pump PUMP CENTRIFUGAL Flow Rate = 1 gpm; Liquid Head = 150 ft 

P‐708 Flue Gas Scrubber Circulation Pump PUMP CENTRIFUGAL Flow = 1,010 GPM; Fluid Head = 60 Ft; Design Pressure = 30 PSIG 

P‐709 Slurry Pump PUMP CENTRIFUGAL Flow = 2 GPM; Fluid Head = 10 Ft; Design Pressure = 45 PSIG 

P‐790 Alcohol Product Pump PUMP CENTRIFUGAL Flow Rate = 100 gpm; Liquid Head = 30 ft 

S‐701 Instrument Air Dryer DRYER PACKAGE UNIT Flow Rate = 400 SCFM; Outlet Dewpoint = ‐40 Deg F 

T‐701 Plant Air Receiver TANK VERTICAL‐VESSEL Vessel Volume = 1,000 Gallons; Design Pressure = 200 PSIG 

T‐702 Firewater Storage Tank TANK FLAT‐BTM‐STORAGE Tank Volume = 600,000 Gallon; Design Pressure = 0.5 PSIG; Flat Roof 

T‐703 Purchased Diesel Storage Tank TANK FLAT‐BTM‐STORAGE Tank Volume = 12,000 Gallon; Design Pressure = 0.5 PSIG; Floating Roof 

T‐704 Ammonia Storage Tank TANK HORIZONTAL‐STORAGE Consumption Rate = 20 Gallon per Day (10 wppm in BFW); 4‐Week Storage Capacity; Tank Volume = 

T‐705 Caustic Storage Tank TANK FLAT‐BTM‐STORAGE Tank Volume = 5,400 Gallon; Design Pressure = 0.5 PSIG; Flat Roof 

T‐706 Olivine / MgO Loading System LOADER PACKAGE UNIT Included in overall cost for gasification & gas clean up based on quotation from Taylor Biomass Energy

T‐707 Tar Reformer Catalyst Loading System LOADER PACKAGE UNIT Solid Catalyst Loading System Including PLC; Maximum Loading Rate of 5 Tons per Day 

T‐708 BFW Chemical Storage Tank TANK HORIZONTAL‐STORAGE Consumption Rate = 14 Gallon per Day (10 wppm in BFW); 4‐Week Storage Capacity; Tank Volume = 

T‐709 Slurry Tank TANK CONE‐BTM‐STORAGE Tank Size = Purge Rate * Residence Time = 1.0 GPM * 2 Days = 2,880 Gallons 

T‐790 Mixed Alcohol Product Storage Tank TANK FLAT‐BTM‐STORAGE Flow Rate = 15 gpm; 7‐Day Storage Capacity; Tank Capacity = 151,200 Gallon; Design Pressure = 5 PS

T‐792 Ethanol Product Storage Tank TANK FLAT‐BTM‐STORAGE Flow Rate = 130 gpm; 7‐Day Storage Capacity; Tank Capacity = 1,310,400 Gallon; Design Pressure = 5 

73
Appendix B. Individual Equipment Cost Summary
 

74
 
Equipment

Number

NumberRequired

NumberSpares

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

Area A100: Feed Handling & Drying

C‐101 4 Hopper Feeder STRM.A100.101 262,454 262,455 1.00 $0 

C‐102 2 Screener Feeder Conveyor STRM.A100.101 262,454 262,455 1.00 $0 

C‐103 2 Radial Stacker Conveyor STRM.A100.101 262,454 262,455 1.00 $0 

C‐104 2 Dryer Feed Screw Conveyor STRM.A100.101 262,454 262,455 1.00 $0 

C‐105 2 Gasifier Feed Screw Conveyor STRM.A100.104 204,131 204,131 1.00 $0 

K‐101 2 Flue Gas Blower STRM.112 977,060 938,445 0.96 $0 

K‐102 2 CO2 Booster Blower STRM.A100.101 262,454 262,455 1.00 $0 

M‐101 4 Hydraulic Truck Dump with Scale STRM.A100.101 262,454 262,455 1.00 $0 

M‐103 3 Front End Loaders STRM.A100.101 262,454 262,455 1.00 $0 

M‐104 2 Waste Heat Dryer STRM.A100.101 262,454 262,455 1.00 $0 S‐101 2 Magnetic Head Pulley STRM.A100.101 262,454 262,455 1.00 $0 

S‐102 2 Vibratory Conveyor STRM.A100.101 262,454 262,455 1.00 $0 

T‐101 4 Dump Hopper STRM.A100.101 262,454 262,455 1.00 $0 

T‐103 2 Dryer Feed Bin STRM.A100.101 262,454 262,455 1.00 $0 

T‐104 2 Dried Biomass Hopper STRM.A100.104 204,131 204,131 1.00 $0 

T‐105 2 Lock Hopper STRM.A100.104 204,131 204,131 1.00 $0 

T‐106 2 Feed Hopper STRM.A100.104 204,131 204,131 1.00 $0 

Area A100 Subtotal 

75
Equipment

Number

NumberRequired

Area A200: Gasification

NumberSpares

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

C‐201 1 Sand / Ash Conditioner / Conveyor STRM.A200.219 2,430 2,428 1.00 $0 

H‐201 1 Syngas Quench Inlet / Combustion Air Pre‐Heat Exchanger * Pinch 1,000 1,000 1.00 $0 H‐202 1 Tar Reformer Effluent / Combustion Air Pre‐Heat Exchanger * Pinch 1,000 1,000 1.00 $0 K‐202 2 Char Combustor Air Blower WORK.A200.A200CC.WK202 8,624 8,624 1.00 $0 M‐201 2 Sand / Ash Cooler STRM.219 2,428 2,428 1.00 $0 

R‐201 2 Indirectly‐Heated Biomass Gasifier STRM.A200.201 500 1,000 2.00 $6,466,667 

R‐202 2 Char Combustor STRM.A200.A200CC.210 5,452,890 5,468,902 1.00 $0 

S‐201 2 Primary Gasifier Cyclone STRM.A200.202 5,237,857 5,236,499 1.00 $0 

S‐202 2 Secondary Gasifier Cyclone STRM.A200.222 254,544 251,092 0.99 $0 

S‐203 2 Primary Combustor Cyclone STRM.A200.A200CC.210 5,452,889 5,468,902 1.00 $0 

S‐204 2 Char Combustor Secondary Cyclone STRM.A200.A200CC.212 502,930 518,943 1.03 $0 

T‐201 1 Sand / Ash Bin STRM.219 2,430 2,428 1.00 $0 Area A200 Subtotal 

76
Equipment

Number

NumberRequired

Area A300: Gas Cleanup

NumberSpares

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

H‐301 1 Quench Water Recirculation Cooler HEAT.A300.A300Q.QCM301 10,000,000 9,917,194 0.99 $0 

H‐302 1 Steam Blowdown / Unreacted Syngas Exchanger Pinch 1,000 1,000 1.00 $65,163 

H‐303 1 Syngas Quench Inlet / Make‐Up Water Exchanger Pinch 1,000 1,000 1.00 $176,012 

H‐304 1 Syngas Quench Inlet / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $35,863 H‐305 1 Alcohol Synthesis Effluent / Combustion Air Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $32,363 

H‐306 1 Syngas Quench Inlet / Combustion Air Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $42,463 H‐307 1 Syngas Quench Inlet / Combustion Air Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $83,563 H‐308 1 Tar Reformer Effluent / Combustion Air Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $26,163 

H‐309 1 Tar Reformer Effluent / Combustion Air Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $59,963 

H‐310 1 Catalyst Regenerator Flue Gas / Tar Reformer Steam & Recycle Exchanger Pinch 1,000 1,000 1.00 $359,712 

H‐311 1 Tar Reformer Effluent / Unreacted Syngas Exchanger Pinch 1,000 1,000 1.00 $47,663 K‐305 1 Combustion Air Blower WORK.A300.A300TR.A300FC.WK305 1,926 1,730 0.90 $2,469,748 M‐300 1 Combustor Fuel Pre‐Mix STRM.A300.A300TR.325A 363,102 379,371 1.04 $0 

M‐300H 1 Tar Reformer Inlet Mix STRM.A300.A300TR.325A 363,102 379,371 1.04 $0 M‐301 1 Syngas Cyclone Separator STRM.A300.A300Q.300 489,600 379,359 0.77 $0 

M‐302 1 Syngas Venturi Scrubber STRM.A300.A300Q.300 489,600 379,359 0.77 $0 

P‐302 1 1 Quench Water Recirculation Pump STRM.A300.A300Q.308 504,844 341,407 0.68 $12,510 

R‐301A 2 Tar Reformer Catalyst Regenerator STRM.A300.225A 500 236,707 473.41 $0 

R‐303 2 Tar Reformer STRM.A300.225A 500 1,000 2.00 $3,233,333 

S‐306 1 Tar Reformer 2‐Stage Cyclone STRM.A300.225A 500 236,707 473.41 $0 S‐307 1 Catalyst Regenerator 2‐Stage Cyclone STRM.A300.225A 500 236,707 473.41 $0 

Area A300 Subtotal 

77
Equipment

Number

NumberRequired

NumberSpares

Area A400: Alcohol Synthesis

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

H‐401 1 Alcohol Synthesis Feed / Effluent Exchanger Pinch 1,000 1,000 1.00 $2,353,181 

H‐402 1 Alcohol Synthesis Feed / Effluent Exchanger Pinch 1,000 1,000 1.00 $2,152,502 

H‐403 1 Compressor Interstage / Deaerator Feed Exchanger Pinch 1,000 1,000 1.00 $450,719 

H‐404 1 Compressor Interstage / Deaerator Feed Exchanger Pinch 1,000 1,000 1.00 $388,090 

H‐405 1 Compressor Interstage / Deaerator Feed Exchanger Pinch 1,000 1,000 1.00 $274,322 

H‐406 1 Compressor Interstage / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $132,222 

H‐407 1 Compressor Interstage / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $67,763 

H‐408 1 Compressor Interstage / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $68,763 

H‐409 1 Alcohol Synthesis Effluent / LO‐CAT Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $22,566 H‐410A 1 1st Stage Air Intercooler HEAT.A400.A400CMPR.QAK‐410A 37,400,000 26,184,715 0.70 $318,900 

H‐410B 1 2nd Stage Air Intercooler HEAT.A400.A400CMPR.QAK‐410B 36,000,000 6,369,531 0.18 $249,600 

H‐410C 1 3rd Stage Air Intercooler HEAT.A400.A400CMPR.QAK‐410C 32,600,000 10,059,981 0.31 $219,300 H‐410D 1 4th Stage Air Intercooler HEAT.A400.A400CMPR.QAK‐410D 38,600,000 6,530,871 0.17 $231,900 H‐410E 1 5th Stage Air Intercooler HEAT.A400.A400CMPR.QAK‐410E 36,000,000 6,001,423 0.17 $244,700 H‐411A 1 1st Stage Water Intercooler HEAT.A400.A400CMPR.QKCW410A 11,100,000 7,941,726 0.72 $128,506 

H‐411B 1 2nd Stage Water Intercooler HEAT.A400.A400CMPR.QKCW410B 6,400,000 4,231,187 0.66 $70,035 

H‐411C 1 3rd Stage Water Intercooler HEAT.A400.A400CMPR.QKCW410C 4,400,000 4,295,205 0.98 $49,835 H‐411D 1 4th Stage Water Intercooler HEAT.A400.A400CMPR.QKCW410D 3,300,000 3,203,051 0.97 $45,735 

H‐411E 1 5th Stage Water Intercooler HEAT.A400.A400CMPR.QKCW410E 3,300,000 3,198,308 0.97 $50,935 

H‐413 1 Syngas Air Cooler HEAT.A400.QAH413GB 33,800,000 26,161,365 0.77 $289,106 

H‐414 1 Syngas Cooling Water Exchanger HEAT.A400.QCH414GB 15,600,000 20,647,374 1.32 $161,435 

H‐415 1 Alcohol Synthesis Effluent / Syngas Recycle Exchanger Pinch 1,000 1,000 1.00 $258,230 

H‐416 1 Alcohol Synthesis Effluent / CO2‐Rich Gas Exchanger Pinch 1,000 1,000 1.00 $28,534 

H‐417 1 Alcohol Synthesis Effluent / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $48,163 

H‐418 1 Alcohol Synthesis Effluent / Expander Inlet Exchanger Pinch 1,000 1,000 1.00 $79,563 

H‐419 1 Tar Reformer Effluent / Expander Inlet Exchanger Pinch 1,000 1,000 1.00 $32,280 

H‐420 1 Tar Reformer Effluent / Expander Inlet Exchanger Pinch 1,000 1,000 1.00 $88,063 

H‐421 1 Alcohol Synthesis Effluent / Expander Inlet Exchanger Pinch 1,000 1,000 1.00 $29,363 

H‐422 1 Tar Reformer Effluent / Expander Inlet Exchanger Pinch 1,000 1,000 1.00 $38,063 H‐423 1 Tar Reformer Effluent / Expander Inlet Exchanger Pinch 1,000 1,000 1.00 $57,963 H‐424 1 Alcohol Synthesis Effluent / CO2‐Rich Gas Exchanger Pinch 1,000 1,000 1.00 $35,963 H‐425 1 Char Combustor Flue Gas / CO2‐Rich Gas Exchanger Pinch 1,000 1,000 1.00 $28,380 

H‐426 1 Tar Reformer Effluent / CO2‐Rich Gas Exchanger Pinch 1,000 1,000 1.00 $41,563 

H‐427 1 Catalyst Regenerator Flue Gas / Alcohol Synthesis Feed Exchanger Pinch 1,000 1,000 1.00 $111,563 

K‐410 1 Synthesis Gas Compressor WORK.A400.WK410 78,374 63,374 0.81 $38,716,300 

K‐414 1 Unconverted Syngas Recycle Compressor WORK.A400.WK414 510 706 1.39 $0 

K‐412A 1 Purge Gas Expander Stage 1 WORK.A400.WK412‐1 ‐10,058 ‐6,342 0.63 $3,300,000 K‐412B 1 Purge Gas Expander Stage 2 WORK.A400.WK412‐2 ‐12,257 ‐9,036 0.74 $9,800,000 

78
Equipment

Number

NumberRequired

NumberSpares

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

M‐401 1 DEPG Acid Gas Removal System STRM.A400.SELEXOL.CO2‐H2S 56,150 60,514 1.08 $6,633,000 M‐402 1 Ammonia Absorption Refrigeration System STRM.A400.SELEXOL.CO2‐H2S 56,150 60,514 1.08 $3,200,000 

M‐403 1 Amine Acid Gas Enrichment System STRM.A400.SELEXOL.CO2‐H2S 56,150 60,514 1.08 $1,241,000 

M‐404 1 Methanol / H2S Absorption System STRM.A400.SELEXOL.RECY‐H2S 6,400 7,749 1.21 $428,000 

M‐405 1 LO‐CAT Sulfur Recovery System STRM.A400.SELEXOL.424 93 39 0.42 $2,917,500 

R‐410 2 Alcohol Synthesis Reactor NTUBES 4,500 4,790 1.06 $5,763,000 

S‐410A 1 1st Interstage KO Drum STRM.313 431,712 351,310 0.81 $72,472 

S‐410B 1 2nd Interstage KO Drum STRM.313 431,712 351,310 0.81 $51,137 

S‐410C 1 3rd Interstage KO Drum STRM.313 431,712 351,310 0.81 $47,835 

S‐410D 1 4th Interstage KO Drum STRM.313 431,712 351,310 0.81 $83,101 

S‐410E 1 5th Interstage KO Drum STRM.313 431,712 351,310 0.81 $108,700 S‐415 1 Pre‐Compressor KO Drum STRM.313 431,712 351,310 0.81 $144,796 

Area A400 Subtotal 

79
Equipment

Number

NumberRequired

NumberSpares

Area A500: Alcohol Separation

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

D‐504 1 Crude Alcohol Distillation Column STRM.A500.507 128,550 122,834 0.96 $1,984,906 

D‐505 1 Methanol Column STRM.A500.A504.510 123,000 116,386 0.95 $3,855,479 

H‐501 1 Ethanol Product / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $30,263 

H‐502 1 Higher Alcohols Product / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $16,684 

H‐503 1 Syngas Quench Inlet / Water‐Methanol Recycle Exchanger Pinch 1,000 1,000 1.00 $244,690 

H‐504C 1 Crude Alcohol Column Condenser HEAT.A500.A504.QAH504 35,939,280 35,157,846 0.98 $370,914 

H‐504R 1 Crude Alcohol Column Reboiler HEAT.A500.A504.QRH504 ‐33,348,297 ‐32,400,695 0.97 $142,353 

H‐505C 1 Methanol Column Condenser HEAT.A500.A504.QAH505 100,301,348 94,492,276 0.94 $949,928 

H‐505R 1 Methanol Column Reboiler HEAT.A500.A504.QRH505 ‐84,068,475 ‐77,564,618 0.92 $299,853 

H‐506 1 Syngas Quench Inlet / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $106,034 

H‐507 1 Alcohol Synthesis Effluent / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $24,080 H‐508 1 Alcohol Synthesis Effluent / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $20,880 H‐509 1 Catalyst Regenerator Flue Gas / Water‐Methanol Recycle Exchanger Pinch 1,000 1,000 1.00 $25,366 H‐510 1 Tar Reformer Effluent / Water‐Methanol Recycle Exchanger Pinch 1,000 1,000 1.00 $24,934 H‐513 1 Molecular Sieve Flush Cooler HEAT.A500.A503MS.QAS513 35,939,280 11,705,592 0.33 $0 

H‐514 1 Methanol Condenser HEAT.A500.A503MS.QAH514 35,939,280 25,880,451 0.72 $370,914 

H‐591 1 Higher Alcohols Product Finishing Cooler HEAT.A500.QCH591 41,542 47,049 1.13 $20,866 

H‐593 1 Ethanol Product Finishing Cooler HEAT.A500.QCH593 381,671 397,956 1.04 $16,780 

P‐504 1 1 Crude Alcohol Column Reflux Pump STRM.A500.A504.510 122,900 116,386 0.95 $6,967 

P‐505 1 1 Methanol Column Reflux Pump STRM.A500.A504.512 73,980 65,385 0.88 $9,571 P‐514 1 1 Condensed Methanol Pump STRM.A500.A503MS.514‐ML1 44,808 44,966 1.00 $396,462 

P‐590 1 1 Crude Alcohol Column Bottoms Pump STRM.A500.A504.511 5,655 6,448 1.14 $8,063 

P‐592 1 1 Methanol Column Bottoms Pump STRM.A500.A504.592A 48,910 51,000 1.04 $11,910 

S‐501 1 Mixed Alcohols Condensation Knock‐out STRM.A400.473A 869,375 1,077,813 1.24 $116,973 

S‐502 1 Mol Sieve Pre‐Flash Drum STRM.A500.503 163,030 162,485 1.00 $60,922 

S‐503 1 Molecular Sieve Separator System STRM.A500.505 66,978 129,421 1.93 $2,600,000 

S‐514 1 Condensed Methanol Flash Drum STRM.A500.A503MS.514‐M0 55,410 55,578 1.00 $19,705 

T‐504 1 Crude Alcohol Column Overhead Accumulator STRM.A500.A504.510 122,900 116,386 0.95 $30,191 

T‐505 1 Methanol Column Overhead Accumulator STRM.A500.A504.512 73,980 65,385 0.88 $46,575 

Area A500 Subtotal 

80
Equipment

Number

NumberRequired

NumberSpares

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

Area A600: Steam System & Power Generation

H‐601 1 Steam Turbine Condenser HEAT.A600.A600ST.QAH601 1.00 5.00 5.00 $1,300,000 

H‐602 1 Steam Turbine Exhaust / Syngas Recycle Exchanger Pinch 1,000 1,000 1.00 $31,463 H‐603 1 Blowdown Cooler HEAT.A600.QCH603 381,671 474,197 1.24 $16,780 

H‐604 1 Steam Turbine Exhaust / Mol‐Sieve Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $30,163 H‐605 1 Syngas Quench Inlet / Make‐Up Water Exchanger Pinch 1,000 1,000 1.00 $27,063 

H‐606 1 Syngas Quench Inlet / BFW Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $107,834 

H‐607 1 Alcohol Synthesis Effluent / BFW Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $164,512 

H‐608 1 Tar Reformer Effluent / BFW Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $726,312 

H‐609 1 Catalyst Regenerator Flue Gas / BFW Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $207,063 

H‐610 1 Char Combustor Flue Gas / BFW Pre‐Heat Exchanger Pinch 1,000 1,000 1.00 $341,534 

H‐611 1 Catalyst Regenerator Flue Gas Steam Generator Pinch 1,000 1,000 1.00 $89,663 H‐612 1 Char Combustor Flue Gas Steam Superheater Pinch 1,000 1,000 1.00 $518,334 

H‐613 1 Tar Reformer Effluent Steam Superheater Pinch 1,000 1,000 1.00 $315,934 

M‐601A 1 BFW Reverse Osmosis (RO) Unit STRM.A600.A600BF.631 300,240 476,788 1.59 $1,865,000 

M‐601B 1 BFW Electrodeionization (EDI) Unit STRM.A600.A600BF.631 300,240 476,788 1.59 $0 M‐601C 1 Hot Condensate Polishing Unit STRM.A600.616+SELSTRET+536 400,439 400,439 1.00 $0 

M‐602A 1 Extraction Steam Turbine / Generator Stage 1 WTURB1 ‐40,418 ‐23,244 0.58 $7,700,000 

M‐602B 1 Extraction Steam Turbine / Generator Stage 2 WTURB2 ‐40,418 ‐47,755 1.18 $7,700,000 

M‐603 1 Startup Boiler STRM.A200.200 36,560 66,542 1.82 $198,351 

M‐604 1 Brine Recovery Reverse Osmosis (RO) Unit STRM.H2O‐MGMT.TOROBRIN 300,240 31,811 0.11 $0 P‐601 1 1 Make‐Up Pump STRM.A600.A600WTR.618 80,411 101,798 1.27 $6,528 

P‐602 1 1 Condensate Pump STRM.A600.633 247,010 476,788 1.93 $9,810 

P‐603 1 1 EDI Pump STRM.A600.633 247,010 476,788 1.93 $9,810 

P‐604 1 1 Boiler Feed Water Pump STRM.A600.633 494,622 476,788 0.96 $304,578 

S‐601 1 Blowdown Flash Drum STRM.A600.604 9,892 9,536 0.96 $47,205 

T‐601 1 Condensate Collection Tank STRM.A600.A600BF.627 500,400 476,788 0.95 $28,505 

T‐602 1 Condensate Surge Tank STRM.A600.A600BF.627 500,400 476,788 0.95 $27,704 

T‐603 1 Deaerator STRM.A600.633 494,619 476,788 0.96 $53,299 

T‐603A 1 Deaerator Packed Column STRM.A600.633 494,619 476,788 0.96 $18,405 T‐605 1 Steam Drum STRM.A600.644 494,622 476,791 0.96 $104,100 

Area A600 Subtotal 

81
Equipment

Number

NumberRequired

NumberSpares

No. Req.Variable Equipment Name Scaling Stream

Scaling

Stream Flow

(lb/hr or btu/hr)

New

Stream

Flow

Size

Ratio

Original

Equip Cost

(per unit)

Area A700: Cooling Water & Other Utilities

K‐701 2 1 Plant Air Compressor STRM.A100.101 262,454 262,455 1.00 $87,922 

M‐701 1 Cooling Tower System STRM.A700.715 7,506,000 5,864,499 0.78 $260,852 M‐702 1 Hydraulic Truck Dump with Scale STRM.A100.101 367,437 262,455 0.71 $80,000 

M‐703 1 Flue Gas Scrubber STRM.112 489,600 938,445 1.92 $436,250 

M‐704 1 Flue Gas Stack STRM.112 939,119 938,445 1.00 $169,187 

M‐705 1 Clarifier 220 220 $96,221 

M‐706 1 Belt Press 1 1 $135,000 

M‐707 1 Sand Filter 265 265 $120,400 

M‐708 1 Emergency Flare STRM.A100.101 262,454 262,455 1.00 $0 

M‐709 1 Cooling Tower Reverse Osmosis (RO) Unit STRM.A100.101 300,240 262,455 0.87 $0 M‐710 1 Product Loading Rack STRM.A100.101 262,454 262,455 1.00 $0 P‐701 1 1 Cooling Water Pump STRM.A700.715 7,001,377 5,864,499 0.84 $239,375 

P‐702 1 1 Firewater Pump STRM.A100.101 262,454 262,455 1.00 $23,043 

P‐703 1 1 Diesel Pump STRM.A100.101 262,454 262,455 1.00 $3,842 

P‐704 1 1 Ammonia Pump STRM.A100.101 262,454 262,455 1.00 $3,842 

P‐705 1 1 Caustic Pump 1 1 $4,906 

P‐707 1 1 BFW Chemical Pump STRM.A100.101 262,454 262,455 1.00 $3,842 

P‐708 1 1 Flue Gas Scrubber Circulation Pump STRM.112 489,600 938,445 1.92 $12,510 

P‐709 1 1 Slurry Pump 1 1 $3,242 P‐790 1 1 Alcohol Product Pump STRM.A100.101 262,454 262,455 1.00 $5,981 

S‐701 1 1 Instrument Air Dryer STRM.A100.101 262,454 262,455 1.00 $8,349 

T‐701 1 Plant Air Receiver STRM.A100.101 262,454 262,455 1.00 $21,005 

T‐702 1 Firewater Storage Tank STRM.A100.101 262,454 262,455 1.00 $229,900 

T‐703 1 Purchased Diesel Storage Tank STRM.A100.101 262,454 262,455 1.00 $104,674 

T‐704 1 Ammonia Storage Tank STRM.A100.101 262,454 262,455 1.00 $15,704 

T‐705 1 Caustic Storage Tank 1 1 $16,005 

T‐706 1 Olivine / MgO Loading System STRM.A100.101 262,454 262,455 1.00 $0 

T‐707 1 Tar Reformer Catalyst Loading System STRM.A100.101 262,454 262,455 1.00 $235,000 

T‐708 1 BFW Chemical Storage Tank STRM.A100.101 262,454 262,455 1.00 $22,004 

T‐709 1 Slurry Tank STRM.A100.101 1 1 1.00 $14,205 

T‐790 1 Mixed Alcohol Product Storage Tank STRM.590 6,063 6,448 1.06 $220,700 

T‐792 1 1 Ethanol Product Storage Tank STRM.592 51,427 51,000 0.99 $690,900 

Area A700 Subtotal 

Plant Total 

82
Appendix C. Discounted Cash Flow Rate of Return (DCFROR) 

and Operating Cost Summary 

83
 
       

    

     

     

             

             

               

               

                       

       

       

         

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Process Engineering Analysis for Ethanol from Mixed Alcohol Synthesis 

2012 Case Using Technical Targets for Tar Reforming & Mixed Alcohol Synthesis
 

2,000 Dry Metric Tonnes Biomass per Day
 

Indirect Gasifier, Tar Reformer, Sulfur Removal, MoS2 Catalyst, Fuel Purification, Steam‐Power Cycle
 

All Values in 2007$
 

Minimum Ethanol Selling Price (MESP) $2.05 per Gallon

Contributions: Feedstock Costs $0.735 per Gallon

Operating Costs & Credits $0.244 per Gallon

Capital Charges & Taxes $1.069 per Gallon

Gasoline‐Equivalent Price $3.11 per Gallon

EtOH Production at Operating Capacity 64.7 MM Gallons per Year

EtOH Product Yield 83.8 Gallons per Dry US Ton Feedstock

Mixed Alcohols Production at Operating Capacity 72.5 MM Gallons per Year

Mixed Alcohols Product Yield 93.9 Gallons per Dry US Ton Feedstock

Delivered Feedstock Cost $61.57 per Dry US Ton (Includes Capital Up to Throat of Gasifier)

Internal Rate of Return (After‐Tax) 10.0%

Equity Percent of Total Investment 40.0%

Capital Costs Operating Costs (¢ / Gallon Product)

Gasification $43,250,000 Feedstock 73.5

Tar Reforming & Quench $26,940,000 Natural Gas 0.0

Acid Gas & Sulfur Removal $28,490,000 Catalysts 9.9

Syngas Compression & Expansion $80,630,000 Olivine & Magnesium Oxide 0.7

Alcohol Synthesis Reaction $41,420,000 Other Raw Materials 0.9

Alcohol Separation $20,310,000 Waste Disposal 0.8

Steam System & Power Generation $45,840,000 Electricity 0.0

Cooling Water & Other Utilities $9,560,000 Fixed Costs 35.9

Total Installed Equipment Cost (TIC) $296,450,000 Co‐Product Credits ‐23.8

Capital Depreciation 37.8

Land (115 acres at $14000 per acre) 1,600,000 Average Income Tax 11.5

Site Development 9,640,000 Average Return on Investment 57.5

(% of ISBL) 4.0%

Indirect Costs & Project Contingency 183,650,000

(% of TIC) 62.0% Operating Costs ($ / Year)

Feedstock $47,560,000

Fixed Capital Investment (FCI) 491,350,000 Natural Gas $0

Working Capital 24,490,000 Catalysts $6,380,000

Total Capital Investment (TCI) 515,840,000 Olivine & Magnesium Oxide $0

Other Raw Materials $380,000

Total Installed Equipment Cost per Annual Gallon Waste Disposal $520,000

of Ethanol Product 4.58 Electricity $0

of Mixed Alcohol Product 4.09 Fixed Costs $23,240,000

Co‐Product Credits at $1.88 per Gallon ‐$15,380,000

Fixed Capital Investment per Annual Gallon Capital Depreciation $24,490,000

of Ethanol Product 7.59 Average Income Tax $7,450,000

of Mixed Alcohol Product 6.78 Average Return on Investment $37,210,000

Loan Interest Rate 8.0% Total Plant Electricity Usage (KW) 64,356

Loan Term (Years) 10 Electricity Produced Onsite (KW) 64,411

Electricity Purchased from Grid (KW) 0

Maximum Yields Based on Feedstock Carbon Content Electricity Sold to Grid (KW) 55

Theoretical Ethanol Production (MM Gal / Year) 159.0

Theoretical Ethanol Yield (Gal / Dry Ton) 205.9 Steam Plant + Turboexpander Power Generated (hp) 86,377

Current Ethanol Yield (Actual / Theoretical) 40.7% Used for Main Compressors (hp) 63,374

Other Electical Consumption (hp) 23,003

Gasifier Efficiency ‐ HHV % 74.3

Gasifier Efficiency ‐ LHV % 73.9 Plant Electricity Use (KWh / Gal EtOH) 8.36

Overall Plant Efficiency ‐ HHV % 46.5 Gasification & Reforming Steam Use (lb / Gal EtOH) 8.68

Overall Plant Efficiency ‐ LHV % 44.9

Specific Operating Conditions

Plant Operating Hours per Year 8,410 Feed Rate Dry Tonnes / Day 2,000

On‐Stream Percentage 96.0% Dry Tons / Day 2,205

Feedstock Cost $ / Dry Ton $61.57

Excel File: R236H‐V29.xls $ / Moisture & Ash Free Ton $62.14

84
Discounted Cash Flow Rate of Return Worksheet 

Year -2 -1 0 1 2 3 4 5 

Fixed Capital Investment $17,281,753 $117,538,146 $62,687,011 

Working Capital $0 $0 $24,487,114 

Loan Payment $0 $0 $0 $43,791,624 $43,791,624 $43,791,624 $43,791,624 $43,791,624 

 Loan Interest Payment $1,880,610 $15,985,188 $23,507,629 $23,507,629 $21,884,910 $20,132,372 $18,239,632 $16,195,473 

 Loan Principal $23,507,629 $199,814,848 $293,845,365 $273,561,369 $251,654,654 $227,995,402 $202,443,410 $174,847,258 $

 Ethanol Sales $115,930,257 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $

 By-Product Credit $13,455,987 $15,378,271 $15,378,271 $15,378,271 $15,378,271 

Total Annual Sales $129,386,244 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $

Annual Manufacturing Cost Raw Materials $44,590,082 $47,562,754 $47,562,754 $47,562,754 $47,562,754 

 Tar reforming catalysts $3,586,607 $0 $0 $0 $0  Mixed Alcohol Catalysts $10,212,364 $0 $8,866,301 $0 $8,866,301 

 Other Variable Costs $3,148,499 $3,323,101 $3,323,101 $3,323,101 $3,323,101 

 Fixed Operating Costs $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 

Total Product Cost $84,776,616 $74,124,919 $82,991,220 $74,124,919 $82,991,220 

Annual Depreciation Total Plant MACRS Schedule 14.29% 24.49% 17.49% 12.49% 8.93% 

Depreciation $69,984,171 $119,937,883 $85,655,924 $61,168,810 $43,733,985 Net Revenue ($48,882,173) ($68,077,719) ($40,909,524) ($5,663,369) $4,949,315 

Losses Forward $0 ($48,882,173) ($116,959,892) ($157,869,415) ($163,532,784) ($Taxable Income ($48,882,173) ($116,959,892) ($157,869,415) ($163,532,784) ($158,583,469) ($Income Tax $0 $0 $0 $0 $0 

Annual Cash Income $818,003 $29,953,449 $21,087,148 $29,953,449 $21,087,148 Discount Factor 1.210 1.100 1.000 0.909 0.826 0.751 0.683 0.621 Annual Present Value $279,248,292 $743,639 $24,754,917 $15,843,086 $20,458,609 $13,093,460 Total Capital Investment + Interest $23,186,459 $146,875,667 $110,681,754 Net Present Worth $0 

9585
Discounted Cash Flow Rate of Return Worksheet 

Year 9 10 11 12 13 14 15 16 17 

Fixed Capital Investment Working Capital Loan Payment $43,791,624 $43,791,624 $0 $0 $0 $0 $0 $0 $0 

 Loan Interest Payment $6,247,365 $3,243,824 $0 $0 $0 $0 $0 $0 $0 

 Loan Principal $40,547,800 $0 $0 $0 $0 $0 $0 $0 $0 

 Ethanol Sales $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $

 By-Product Credit $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 

Total Annual Sales $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $

Annual Manufacturing Cost Raw Materials $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 

 Tar reforming catalysts $0 $0 $0 $0 $0 $0 $0 $0 $0 

 Mixed Alcohol Catalysts $8,866,301 $0 $8,866,301 $0 $8,866,301 $0 $8,866,301 $0 $8,866,301 

 Other Variable Costs $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 

 Fixed Operating Costs $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 

Total Product Cost $82,991,220 $74,124,919 $82,991,220 $74,124,919 $82,991,220 $74,124,919 $82,991,220 $74,124,919 $82,991,220 

Annual Depreciation Total Plant MACRS Schedule 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 

Depreciation $0 $0 $0 $0 $0 $0 $0 $0 $0 

Net Revenue $58,631,408 $70,501,250 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 

Losses Forward ($90,095,725) ($31,464,318) $0 $0 $0 $0 $0 $0 $0 

Taxable Income ($31,464,318) $39,036,932 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 

Income Tax $0 $13,662,926 $22,707,570 $25,810,776 $22,707,570 $25,810,776 $22,707,570 $25,810,776 $22,707,570 

Annual Cash Income $21,087,148 $16,290,523 $42,171,202 $47,934,298 $42,171,202 $47,934,298 $42,171,202 $47,934,298 $42,171,202 

Discount Factor 0.424 0.386 0.350 0.319 0.290 0.263 0.239 0.218 0.198 

Annual Present Value $8,943,009 $6,280,702 $14,780,749 $15,273,345 $12,215,495 $12,622,599 $10,095,451 $10,431,900 $8,343,348 Total Capital Investment + Interest Net Present Worth 

9686
Discounted Cash Flow Rate of Return Worksheet 

Year 21 22 23 24 25 26 27 28 29 

Fixed Capital Investment Working Capital (Loan Payment $0 $0 $0 $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 

 Ethanol Sales $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $132,491,722 $

 By-Product Credit $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 $15,378,271 

Total Annual Sales $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $147,869,993 $

Annual Manufacturing Cost Raw Materials $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 $47,562,754 

 Tar reforming catalysts $0 $0 $0 $0 $0 $0 $0 $0 $0 

 Mixed Alcohol Catalysts $8,866,301 $0 $8,866,301 $0 $8,866,301 $0 $8,866,301 $0 $8,866,301 

 Other Variable Costs $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 $3,323,101 

 Fixed Operating Costs $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 $23,239,064 

Total Product Cost $82,991,220 $74,124,919 $82,991,220 $74,124,919 $82,991,220 $74,124,919 $82,991,220 $74,124,919 $82,991,220 Annual Depreciation Total Plant MACRS Schedule 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 0.00% 

Depreciation $0 $0 $0 $0 $0 $0 $0 $0 $0 

Net Revenue $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 Losses Forward $0 $0 $0 $0 $0 $0 $0 $0 $0 

Taxable Income $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 $73,745,074 $64,878,773 

Income Tax $22,707,570 $25,810,776 $22,707,570 $25,810,776 $22,707,570 $25,810,776 $22,707,570 $25,810,776 $22,707,570 Annual Cash Income $42,171,202 $47,934,298 $42,171,202 $47,934,298 $42,171,202 $47,934,298 $42,171,202 $47,934,298 $42,171,202 Discount Factor 0.135 0.123 0.112 0.102 0.092 0.084 0.076 0.069 0.063 

Annual Present Value $5,698,619 $5,888,535 $4,709,602 $4,866,558 $3,892,233 $4,021,949 $3,216,722 $3,323,925 $2,658,448 Total Capital Investment + Interest Net Present Worth 

9787
Base Case Operating Cost Summary

 Costing Code Raw Material Stream Number kg / Hr (or kW) lb / Hr (or HP) Quoted Price (Cents / Ton or Cents / kWh) Year of Price Quote 2007 Cost (Cents / Ton) 2007 Cost ($ / lb) 

Variable Operating Costs

Feedstock, Catalyst & Chemicals

 F-3 Feedstock - Wood Chips STRM.100 119,048 262,455 4,310 2007 4,310 0.022  P-2 Electricity WORK.WNET (47.64) (63.89) 5.85 2007 

 R-12 Magnesium Oxide (MgO) STRM.220 3.2 7.0 36,500 2004 42,942 0.215  R-13 Fresh Olivine STRM.221 243.8 537.6 17,290 2004 20,342 0.102 

 R-14 Tar Reformer Catalyst STRM.A300.A300TR.326 5.4 12.0 3,528,850 2007 3,528,850 17.644 

 S-20 50 wt% Caustic Calc'd on SO2 in Flue Gas 18.1 40.0 15,000 2010 13,038 0.065 

 S-21 Boiler Chemicals STRM.635 1.5 3.3 280,000 1991 453,217 2.266 

 S-22 Cooling Tower Chemicals STRM.711 0.5 1.2 200,000 1999 271,610 1.358 

 S-23 Natural Gas for Reformer 0.0 0.1 30,177 2007 30,177 0.151 

 S-24 Cooling Tower Makeup STRM.710 41,311 91,076 19.96 2001 26 0.000  S-25 Boiler Feed Water Makeup STRM.618 34,631 76,348 19.96 2001 26 0.000 

 S-26 Natural Gas for Gasifier STRM.206 0.0 0.0 30,177 2007 30,177 0.151 

 S-27 Diesel Fuel Assumed rate of 10 gal / hr 32.2 70.9 80,589 2007 80,589 0.403 

 S-28 LO-CAT Chemicals STRM.A400.SELEXOL.424 160,150 2009 Annual cost scaled on sulfur produc

 S-31 DEPG Make-Up STRM.A400.SELEXOL.CO2-H2S 26,702 2004 Annual cost scaled on CO2 + H2S re S-32 Amine Make-Up STRM.A400.SELEXOL.CO2-H2S 8,000 2010 Annual cost scaled on CO2 + H2S re Subtotal 

Waste Streams

 W-1 Sand and Ash Purge STRM.219 1,101 2,428 3,266 1998 4,465 0.022 

 W-2 Tar Reformer Catalyst Disposal 5.4 12.0 3,266 1998 4,465 0.022 

 W-3 WWT Cost STRM.305 13,129 28,945 48.08 2001 62 0.000 

 Subtotal 

Co-Product Credits

 Mixed Alcohol STRM.590 2,925 6,448 56,722 2007 56,722 0.284 

 Subtotal 

Total Variable Operating Costs 

Fixed Operating Costs

 2007 Salary # Positions 2007 Cost 2007 Hourly Wage

 Plant Manager $147,000 1 $147,000 $70.67 Plant Engineer $70,000 1 $70,000 $33.65 Maintenance Supr $57,000 1 $57,000 $27.40 Lab Manager $56,000 1 $56,000 $26.92 Shift Supervisor $48,000 5 $240,000 $23.08 Lab Technician $40,000 2 $80,000 $19.23 Maintenance Tech $40,000 16 $640,000 $19.23

 Shift Operators $48,000 20 $960,000 $23.08 Yard Employees $28,000 12 $336,000 $13.46

 Clerks & Secretaries $36,000 3 $108,000 $17.31

 Total Salaries 62 $2,694,000 

Avg Salary (w/ Benefits)

 Overhead and Benefits % of Labor & Supervison = 90.0% $2,424,600 $39.69 per Hour  Maintenance % of FCI = 3.0% $14,694,771 9,184,232 $82,558 per Year  Insurance & Taxes % of FCI = 0.7% $3,428,780 3,428,780 

Total Fixed Operating Costs 

Total Operating Costs 

88
Optimized Water Consumption Case Operating Cost Summary

 Costing Code Raw Material Stream Number kg / Hr (or kW) lb / Hr (or HP) Quoted Price (Cents / Ton or Cents / kWh) Year of Price Quote 2007 Cost (Cents / Ton) 2007 Cost ($ / lb) 

Variable Operating Costs

Feedstock, Catalyst & Chemicals

 F-3 Feedstock - Wood Chips STRM.100 119,048 262,455 4,310 2007 4,310 0.022  P-2 Electricity WORK.WNET (55.36) (74.24) 5.85 2007 

 R-12 Magnesium Oxide (MgO) STRM.220 3.2 7.0 36,500 2004 42,942 0.215  R-13 Fresh Olivine STRM.221 243.8 537.6 17,290 2004 20,342 0.102 

 R-14 Tar Reformer Catalyst STRM.A300.A300TR.326 5.4 12.0 3,528,850 2007 3,528,850 17.644 

 S-20 50 wt% Caustic Calc'd on SO2 in Flue Gas 18.1 40.0 15,000 2010 13,038 0.065 

 S-21 Boiler Chemicals STRM.635 1.5 3.3 280,000 1991 453,217 2.266 

 S-22 Cooling Tower Chemicals STRM.711 0.5 1.2 200,000 1999 271,610 1.358 

 S-23 Natural Gas for Reformer 0.0 0.1 30,177 2007 30,177 0.151 

 S-24 Cooling Tower Makeup STRM.710 10,918 24,070 19.96 2001 26 0.000  S-25 Boiler Feed Water Makeup STRM.618 46,175 101,798 19.96 2001 26 0.000 

 S-26 Natural Gas for Gasifier STRM.206 0.0 0.0 30,177 2007 30,177 0.151 

 S-27 Diesel Fuel Assumed rate of 10 gal / hr 32.2 70.9 80,589 2007 80,589 0.403 

 S-28 LO-CAT Chemicals STRM.A400.SELEXOL.424 160,150 2009 Annual cost scaled on sulfur produc

 S-31 DEPG Make-Up STRM.A400.SELEXOL.CO2-H2S 26,702 2004 Annual cost scaled on CO2 + H2S re S-32 Amine Make-Up STRM.A400.SELEXOL.CO2-H2S 8,000 2010 Annual cost scaled on CO2 + H2S re Subtotal 

Waste Streams

 W-1 Sand and Ash Purge STRM.219 1,101 2,428 3,266 1998 4,465 0.022 

 W-1A Scrubber Solids STRM.H2O-MGMT.SCRB-SLD 101 224 3,266 1998 4,465 0.022  W-2 Tar Reformer Catalyst Disposal 5.4 12.0 3,266 1998 4,465 0.022 

 W-3 WWT Cost STRM.H2O-MGMT.ROBRNWST 3,607 7,953 48.08 2001 62 0.000  Subtotal 

Co-Product Credits

 Mixed Alcohol STRM.590 2,925 6,448 56,722 2007 56,722 0.284 

 Subtotal 

Total Variable Operating Costs 

Fixed Operating Costs

 2007 Salary # Positions 2007 Cost 2007 Hourly Wage

 Plant Manager $147,000 1 $147,000 $70.67 Plant Engineer $70,000 1 $70,000 $33.65 Maintenance Supr $57,000 1 $57,000 $27.40 Lab Manager $56,000 1 $56,000 $26.92 Shift Supervisor $48,000 5 $240,000 $23.08 Lab Technician $40,000 2 $80,000 $19.23 Maintenance Tech $40,000 16 $640,000 $19.23

 Shift Operators $48,000 20 $960,000 $23.08 Yard Employees $28,000 12 $336,000 $13.46

 Clerks & Secretaries $36,000 3 $108,000 $17.31

 Total Salaries 62 $2,694,000 

Avg Salary (w/ Benefits)

 Overhead and Benefits % of Labor & Supervison = 90.0% $2,424,600 $39.69 per Hour  Maintenance % of FCI = 3.0% $14,692,268 9,182,668 $82,558 per Year  Insurance & Taxes % of FCI = 0.7% $3,428,196 3,428,196 

Total Fixed Operating Costs 

Total Operating Costs 

89
Appendix D. Process Parameters and Operation Summary
 

90
 
Base Case Process Parameters and Operation Summary

Energy Efficiencies Tar Reformer Alcohol Synthesis Alcohol Synthesis AlcoGasifier Efficiency - HHV % 74.3 Inlet Molar Flow (MMscf/hr) 6.72 Syngas from Conditioning 315,330 Relative Alcohol Distribution After Reactor Upstr

Gasifier Efficiency - LHV % 73.9 Space Velocity (hr-1) 2,476 Unreacted Syngas Recycle 709,768 Methanol 61.6% Tem

Overall Plant Efficiency - HHV % 46.5 Reactor Volume (ft3) 2,713 Recycled Methanol from Mol Sieve 44,966 Ethanol 35.6% PreOverall Plant Efficiency - LHV % 44.9 Catalyst Replacement (%inventory/day) 0.10% Recycled Methanol and H2S 7,749 Propanol 2.8%

Inlet: Total 1,077,813 Butanol 0.0% LP S

Dryer Temperature (°F) 1,632 Pentanol + 0.0% Tem

Inlet: Carbon as CO (mol%) 31.8% Conditioned Syngas H2:CO Ratio 1.25 Pre Temperature (°F) 60.0 Carbon as tar (ppmv) 8,089 Recycled Gas H2:CO Ratio 1.67 Flash Separator Moisture Content (wt%) 30.0% H2:CO Ratio (mole) 0.86 Temperature (°F) 110 LP S

Outlet: Reformer Conversions: At Reactor Inlet Pressure (psia) 2,937 Me

 Temperature (°F) 219.7 CO2 --> CO 24.6% Temperature (°F) 595 Eth

 Moisture Content (wt%) 10.0% Methane --> CO 80.0% Pressure (psia) 2,995 Relative Alcohol Distribution After Flash Tank Pro

Inlet Flue Gas (°F) 539 Ethane --> CO 99.0% H2:CO Molar Ratio 1.50 Methanol 61.3% ButOutlet Flue Gas (°F) 167.9 Ethylene --> CO 90.0% CO2 (mol %) 14.0% Ethanol 35.8% Pen

 Dew Point Flue Gas (°F) 152.8 Benzene --> CO 100.0% Methane (mol%) 7.7% Propanol 2.8% Wa Difference 15.1 Tar --> CO 99.5% H2O (wt%) 0.04% Butanol 0.0%

 Ammonia --> CO 100.0% Methanol (mol%) 2.6% Pentanol + 0.0% Mole 

Gasifier Inlet Molar Flow (MMscf/hr) 23.2 H2Temperature (°F) 1,596 Outlet: Space Velocity (hr-1) 5,000 Vapor Losses From Flash Tank Me

Pressure (psia) 33.0 Temperature (°F) 1,670 Reactor Volume (ft3) 4,365 Methanol 2.1% Eth

H2:CO Molar Ratio After Gasifier 0.56 Carbon as CO (mol%) 73.7% Reactor Length (feet) 60 Ethanol 1.0% Pro

Methane (vol%) 8.3% Carbon as tar (ppmv) 82.1 No of Tubes 9,579 Propanol 0.4% But

Benzene (vol%) 0.07% H2:CO Ratio (mole) 1.25 Catalyst Density (lb/ft3) 75 Butanol 0.1% Pen

Tar (wt%) 0.89% Methane (vol%) 1.5% Pentanol +Tar (g/Nm3) 9.6 Benzene (ppmv) 0.0 CO Conversion - Overall 78.5% RelatChar (wt%) 14.3% Tars (ppmv) 3.6 CO Conversion - Singlepass 28.7% Cleaned Gas Recycled to Reactor 88.5% Me

H2S (ppm) 138.7 Tars (g/Nm3) 0.02 Conversion To: Eth

Residual Heat (MBtu/hr) 29,379 H2S (ppm) 92.6 CO2 8.9% Sulfur (H2S) at Reactor Inlet (spec) ppmv 70.0 Pro

Raw Dry Syngas Yield (lb/lb dry feed) 0.78 NH3 (ppm) 0.0 Methane 3.5% But

 Methane (mole% dry basis) 1.8% Ethane 0.2% Pe

Char Combustor Quench Methanol 2.0% Residual SyngasTemperature (°F) 1,808 Benzene (ppmv) 3.0 Ethanol 12.3% Recycled to synthesis reactors (lb/hr) 709,768 Ethan

Pressure (psia) 29.0 Tars (ppmv) 3.2 Propanol 1.5% To Tar Reformers (lb/hr) 145,046 Init

Ratio Actual:Minimum air for combustion 1.20 Tars (g/Nm3) 0.01 Butanol 0.0% To Fuel System (lb/hr) 36,261 LP 

Residual Heat (MBtu/hr) -0.1 H2S (ppm) 114.2 Pentanol + 0.0% To Reformer for Process (lb/hr) 108,784 C2

NH3 (ppm) 57.1 Total 28.5% Fin

Syngas Usage OvTo Reformer (lb/hr) 236,707 Acid Gas Removal Selectivity (CO2 Free)

To Char Combustor (lb/hr) 0 Inlet: Alcohols 80.8% Electricity

To Fuel System (lb/hr) 9,080 CO2 (mol/hr) 9,386 Hydrocarbons 19.2% Electricity Produced (kWh) 541,655,443

 CO2 (mol%) 19.1% Ethanol 63.1% Electricity Produced (kWh/gallon EtOH) 8.37

Fuel System H2S (mol/hr) 2.9 At Reactor Outlet Electricity Used (kWh) 541,254,810

Additional fuel (lb/hr) 0 H2S (ppmv) 59.0 Temperature (°F) 611 Electricity Used (kWh/gallon) 8.36

Raw Syngas (lb/hr) 9,080 Outlet: Pressure (psia) 2,953

Unconverted Syngas (lb/hr) 36,261 CO2 (mol/hr) 7,243 CO2 (mol%) 18.9% Water Usage

 CO2 (mol%) 15.7% Methane (mol%) 10.0% Boiler Feed Water Makeup (MM lb/yr) 642

Into Reformer (°F) 3,403 Fraction CO2 removed 22.8% H2O (wt%) 0.67% Cooling Water Makeup (MM lb/yr) 766Out of Reformer (°F) 1,777 H2S (mol/hr) 0.3 Total Makeup Water (MM gal/yr) 169

 H2S (ppmv) 6.0 Total Alcohol Productivity (kg/kg/hr) 0.368 Water Usage (gal/gal ETOH) 2.61 Fraction H2S removed 99.5% Total Ethanol Productivity (kg/kg/hr) 0.160

91
Optimized Water Consumption Case Process Parameters and Operation Summary

Energy Efficiencies Tar Reformer Alcohol Synthesis Alcohol Synthesis AlcoGasifier Efficiency - HHV % 74.3 Inlet Molar Flow (MMscf/hr) 6.72 Syngas from Conditioning 315,330 Relative Alcohol Distribution After Reactor Upstr

Gasifier Efficiency - LHV % 73.9 Space Velocity (hr-1) 2,476 Unreacted Syngas Recycle 709,768 Methanol 61.6% Tem

Overall Plant Efficiency - HHV % 46.5 Reactor Volume (ft3) 2,713 Recycled Methanol from Mol Sieve 44,966 Ethanol 35.6% PreOverall Plant Efficiency - LHV % 44.9 Catalyst Replacement (%inventory/day) 0.10% Recycled Methanol and H2S 7,749 Propanol 2.8%

Inlet: Total 1,077,813 Butanol 0.0% LP S

Dryer Temperature (°F) 1,632 Pentanol + 0.0% Tem

Inlet: Carbon as CO (mol%) 31.8% Conditioned Syngas H2:CO Ratio 1.25 Pre Temperature (°F) 60.0 Carbon as tar (ppmv) 8,089 Recycled Gas H2:CO Ratio 1.67 Flash Separator Moisture Content (wt%) 30.0% H2:CO Ratio (mole) 0.86 Temperature (°F) 110 LP S

Outlet: Reformer Conversions: At Reactor Inlet Pressure (psia) 2,937 Me

 Temperature (°F) 219.7 CO2 --> CO 24.6% Temperature (°F) 595 Eth

 Moisture Content (wt%) 10.0% Methane --> CO 80.0% Pressure (psia) 2,995 Relative Alcohol Distribution After Flash Tank Pro

Inlet Flue Gas (°F) 539 Ethane --> CO 99.0% H2:CO Molar Ratio 1.50 Methanol 61.3% ButOutlet Flue Gas (°F) 167.9 Ethylene --> CO 90.0% CO2 (mol %) 14.0% Ethanol 35.8% Pen

 Dew Point Flue Gas (°F) 152.8 Benzene --> CO 100.0% Methane (mol%) 7.7% Propanol 2.8% Wa Difference 15.1 Tar --> CO 99.5% H2O (wt%) 0.04% Butanol 0.0%

 Ammonia --> CO 100.0% Methanol (mol%) 2.6% Pentanol + 0.0% Mole 

Gasifier Inlet Molar Flow (MMscf/hr) 23.2 H2Temperature (°F) 1,596 Outlet: Space Velocity (hr-1) 5,000 Vapor Losses From Flash Tank Me

Pressure (psia) 33.0 Temperature (°F) 1,670 Reactor Volume (ft3) 4,365 Methanol 2.1% Eth

H2:CO Molar Ratio After Gasifier 0.56 Carbon as CO (mol%) 73.7% Reactor Length (feet) 60 Ethanol 1.0% Pro

Methane (vol%) 8.3% Carbon as tar (ppmv) 82.1 No of Tubes 9,579 Propanol 0.4% But

Benzene (vol%) 0.07% H2:CO Ratio (mole) 1.25 Catalyst Density (lb/ft3) 75 Butanol 0.1% Pen

Tar (wt%) 0.89% Methane (vol%) 1.5% Pentanol +Tar (g/Nm3) 9.6 Benzene (ppmv) 0.0 CO Conversion - Overall 78.5% RelatChar (wt%) 14.3% Tars (ppmv) 3.6 CO Conversion - Singlepass 28.7% Cleaned Gas Recycled to Reactor 88.5% Me

H2S (ppm) 138.7 Tars (g/Nm3) 0.02 Conversion To: Eth

Residual Heat (MBtu/hr) 29,379 H2S (ppm) 92.6 CO2 8.9% Sulfur (H2S) at Reactor Inlet (spec) ppmv 70.0 Pro

Raw Dry Syngas Yield (lb/lb dry feed) 0.78 NH3 (ppm) 0.0 Methane 3.5% But

 Methane (mole% dry basis) 1.8% Ethane 0.2% Pe

Char Combustor Quench Methanol 2.0% Residual SyngasTemperature (°F) 1,808 Benzene (ppmv) 3.0 Ethanol 12.3% Recycled to synthesis reactors (lb/hr) 709,768 Ethan

Pressure (psia) 29.0 Tars (ppmv) 3.2 Propanol 1.5% To Tar Reformers (lb/hr) 145,046 Init

Ratio Actual:Minimum air for combustion 1.20 Tars (g/Nm3) 0.01 Butanol 0.0% To Fuel System (lb/hr) 36,261 LP 

Residual Heat (MBtu/hr) -0.1 H2S (ppm) 114.2 Pentanol + 0.0% To Reformer for Process (lb/hr) 108,784 C2

NH3 (ppm) 57.1 Total 28.5% Fin

Syngas Usage OvTo Reformer (lb/hr) 236,707 Acid Gas Removal Selectivity (CO2 Free)

To Char Combustor (lb/hr) 0 Inlet: Alcohols 80.8% Electricity

To Fuel System (lb/hr) 9,080 CO2 (mol/hr) 9,386 Hydrocarbons 19.2% Electricity Produced (kWh) 541,672,459

 CO2 (mol%) 19.1% Ethanol 63.1% Electricity Produced (kWh/gallon EtOH) 8.37

Fuel System H2S (mol/hr) 2.9 At Reactor Outlet Electricity Used (kWh) 541,206,899

Additional fuel (lb/hr) 0 H2S (ppmv) 59.0 Temperature (°F) 611 Electricity Used (kWh/gallon) 8.36

Raw Syngas (lb/hr) 9,080 Outlet: Pressure (psia) 2,953

Unconverted Syngas (lb/hr) 36,261 CO2 (mol/hr) 7,243 CO2 (mol%) 18.9% Water Usage

 CO2 (mol%) 15.7% Methane (mol%) 10.0% Boiler Feed Water Makeup (MM lb/yr) 856

Into Reformer (°F) 3,403 Fraction CO2 removed 22.8% H2O (wt%) 0.67% Cooling Water Makeup (MM lb/yr) 202Out of Reformer (°F) 1,777 H2S (mol/hr) 0.3 Total Makeup Water (MM gal/yr) 127

 H2S (ppmv) 6.0 Total Alcohol Productivity (kg/kg/hr) 0.368 Water Usage (gal/gal ETOH) 1.97 Fraction H2S removed 99.5% Total Ethanol Productivity (kg/kg/hr) 0.160

92
Appendix E. Process Flow Diagrams (PFDs) and Material and 

Energy Balances 

93
 
Heat Stream No. MM BTU/hr Work Stream No. HP

COMPONENT UNITS 101

Total Flow l b/hr 262,455

Temp eratureF 60

Press ure Psia 14.7

Vapor Fraction 0.00

Hydrogen l b/hr-

Carbon Monoxide l b/hr-

Nitrogen l b/hr-

Oxyge n l b/hr-

Argon l b/hr-

Carbon Di oxide l b/hr-

Wate r l b/hr 78,736

Hydrogen Sul fi de (H2S) l b/hr-

SO2 l b/hr-

Ammo nia (NH3) l b/hr-

NO2 l b/hr-

Metha ne (CH4) l b/hr-

Ethane (C2H6) l b/hr-

Ethyle ne (C2H4) l b/hr-

Acetyl ene (C2H2) l b/hr-

Propa ne (C3H8) l b/hr-

I-Buta ne (C4H10) l b/hr-

N-Butane (C4H10) l b/hr-

Penta ne + l b/hr-

Benzene (C6H6) l b/hr-

Tar (C10H8) l b/hr-

Carbon (Solid) l b/hr-

Sulfur (Sol id) l b/hr-

Olivine + MgO (Soli d) l b/hr-

Metha nol l b/hr-

Ethanol l b/hr-

Propa nol l b/hr-

n-Butanol l b/hr-

Penta nol+ l b/hr-

Ash l b/hr0

Char l b/hr0

Wood l b/hr 183,718

Entha lpy Flow MMBTU/hr (972.9)

94
Heat Stream No. MM BTU/hr Work Stream No. HP

QM104 (86.73)

COMPONENT UNITS 101 102 103 104 105 107 110 CO2VENT

Total Flow lb/hr 262,455 262,455 204,131 204,131 204,131 878,723 937,046 1,062

TemperatureF 60 60 220 220 220 539 168 149

Press ure Psia 14.7 15.7 15.7 15.7 35.0 17.0 15.7 18.0

Vapor Fra ction 0.00 0.00 0.00 0.00 0.00 1.00 1.00 1.00

Hydrogen lb/hr - - - - - - - -

Carbon Monoxide lb/hr - - - - - - - -

Nitrogen lb/hr - - - - - 541,810 541,810 -

Oxyge n lb/hr - - - - - 25,738 25,738 -

Argon lb/hr - - - - - 9,236 9,236 -

Carbon Di oxide lb/hr - - - - - 231,742 231,742 1,028

Water lb/hr 78,736 78,736 20,413 20,413 20,413 70,097 128,420 34

Hydrogen Sulfide (H2S) lb/hr-------0

SO2 lb/hr - - - - - 29 29 -

Ammonia (NH3) lb/hr - - - - - - - -

NO2 lb/hr - - - - - 69 69 -

Methane (CH4) lb/hr - - - - - - - -

Ethane (C2H6) lb/hr - - - - - - - -

Ethyle ne (C2H4) lb/hr - - - - - - - -

Acetylene (C2H2) lb/hr - - - - - - - -

Propa ne (C3H8) lb/hr - - - - - - - -

I-Buta ne (C4H10) lb/hr - - - - - - - -

N-Butane (C4H10) lb/hr - - - - - - - -

Penta ne + lb/hr - - - - - - - -

Benzene (C6H6) lb/hr - - - - - - - -

Tar (C10H8) lb/hr - - - - - - - -

Carbon (Solid) lb/hr - - - - - - - -

Sulfur (Solid) lb/hr - - - - - - - -

Olivine + MgO (Sol id) lb/hr - - - - - 2 2 -

Methanol lb/hr - - - - - - - -

Ethanol lb/hr - - - - - - - -

Propa nol lb/hr - - - - - - - -

n-Butanol lb/hr - - - - - - - -

Penta nol+ lb/hr - - - - - - - -

Ash lb/hr 0 0 0 0 0 - - -

Char lb/hr 0 0 0 0 0 0 0 -

Wood lb/hr 183,718 183,718 183,718 183,718 183,718 ---

Entha lpy Flow MMBTU/hr (972.9) (972.9) (553.0) (553.0) (553.0) (1,191.0) (1,610.9) (4.1)

95
Heat Stream No. MM BTU/hr Work Stream No. HP

QH201 4.64 WK202 8,624.1

QH202 9.95

COMPONENT UNITS 105 200 200METH2 202 203 208 209A 209B 210 211 212 220 221 222 223 224 225A 225F 513L-G1

Total Flow l b/hr 204,131 66,542 82,409 5,236,499 4,985,407 477,646 477,646 477,646 5,468,902 4,949,959 518,943 7 538 251,092 5,305 245,787 236,707 9,080 15,867

TemperatureF 220 590 590 1,596 1,596 90 277 400 1,808 1,808 1,808 60 60 1,596 1,596 1,596 1,596 1,596 590

Press ure Psi a 35.0 35.0 35.0 33.0 31.8 14.7 31.0 30.0 29.0 35.0 27.0 31.0 31.0 31.8 31.4 31.4 31.4 31.4 35.0

Vapor Fra ction 0.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 1.00 0.00 1.00 1.00 1.00 1.00

Hydrogen l b/hr--0 3,076 -------- - 3,076 - 3,076 2,963 114 0

Carbon Monoxide l b/hr--0 75,802 -------- - 75,802 - 75,802 73,002 2,800 0

Nitrogen l b/hr-- 0 0 - 353,688 353,688 353,688 353,709 - 353,709 -- 0 - 0 0 0 0

Oxygen l b/hr----- 108,371 108,371 108,371 18,058 - 18,058 --------

Argon l b/hr----- 6,032 6,032 6,032 6,032 - 6,032 --------

Carbon Di oxide l b/hr-- 69 36,225 - 235 235 235 107,159 - 107,159 - - 36,225 - 36,225 34,886 1,338 69

Water l b/hr 20,413 66,542 73,120 93,533 - 9,320 9,320 9,320 31,792 - 31,792 - - 93,533 - 93,533 90,078 3,455 6,578

Hydrogen Sulfide (H2S) l b/hr--2 56 ------- 0 0 56 - 56 54 2 2

SO2 l b/hr----- - -- 9 - 9 -- - -----

Ammonia (NH3) l b/hr--2 356 --------- 356 - 356 343 13 2

NO2 l b/hr----- - ------- - -----

Methane (CH4) l b/hr--0 15,816 -------- - 15,816 - 15,816 15,232 584 0

Ethane (C2H6) l b/hr--0 570 --------- 570 - 570 549 21 0

Ethylene (C2H4) l b/hr--0 7,840 -------- - 7,840 - 7,840 7,550 290 0

Acetylene (C2H2) l b/hr--0 679 --------- 679 - 679 654 25 0

Propa ne (C3H8) l b/hr-- 24 24 --------- 24 - 24 23 1 24

I-Buta ne (C4H10) l b/hr----- - ------- - -----

N-Butane (C4H10) l b/hr-- 0 0 - - ------- 0 - 0 0 0 0

Penta ne + l b/hr----- - ------- - -----

Benzene (C6H6) l b/hr--0 641 --------- 641 - 641 617 24 0

Tar (C10H8) l b/hr--- 1,923 -------- - 1,923 - 1,923 1,852 71 -

Carbon (Soli d) l b/hr----- - ------- - -----

Sulfur (Solid) l b/hr----- - ------- - -----

Olivine + MgO (Sol id) l b/hr--- 4,949,870 4,944,920 --- 4,950,365 4,949,870 495 7 538 4,950 4,900 49 48 2 -

Methanol l b/hr-- 9,069 9,069 -------- - 9,069 - 9,069 8,734 335 9,069

Ethanol l b/hr-- 76 76 --------- 76 - 76 73 3 76

Propa nol l b/hr----- - ------- - -----

n-Butanol l b/hr----- - ------- - -----

Penta nol+ l b/hr----- - ------- - -----

Ash l b/hr 0 -- 0 0 - -- 1,779 89 1,690 -- 0 0 0 0 0 -

Char l b/hr 0 -- 40,896 40,487 -------- 409 405 4 4 0 -

Wood l b/hr 183,718 -- 0 0 -------- 0 0 0 0 0 -

Entha lpy Flow MMBTU/hr (553.0) (368.7) (428.3) 1,196.8 1,877.9 (53.3) (31.3) (16.7) 1,862.9 2,207.5 (344.7) (0.0) (0.0) (681.0) 1.7 (682.8)

96

(657.5) (25.2) (59.6)
Heat Stream No. MM BTU/hr Work Stream No. HP

QCM201 0.86

COMPONENT UNITS 107 212 216 218 219 286 287 334

Total Flow lb/hr 878,723 518,943 2,185 243 2,428 516,758 516,758 361,965

TemperatureF 539 1,808 1,808 66 206 1,808 532 549

Press ure Psia 17.0 27.0 26.7 75.0 14.7 27.0 23.0 17.0

Vapor Fra ction 1.00 1.00 0.00 0.00 0.00 1.00 1.00 1.00

Hydrogen lb/hr - - - - - - - -

Carbon Monoxide lb/hr - - - - - - - -

Nitrogen lb/hr 541,810 353,709 --- 353,709 353,709 188,101

Oxyge n lb/hr 25,738 18,058 --- 18,058 18,058 7,680

Argon lb/hr 9,236 6,032 --- 6,032 6,032 3,204

Carbon Di oxide lb/hr 231,742 107,159 --- 107,159 107,159 124,583

Wate r lb/hr 70,097 31,792 - 243 243 31,792 31,792 38,305

Hydrogen Sulfide (H2S) lb/hr--------

SO2 lb/hr 29 9 --- 9 9 20

Ammonia (NH3) lb/hr - - - - - - - -

NO2 lb/hr 69 ------ 69

Methane (CH4) lb/hr - - - - - - - -

Ethane (C2H6) lb/hr - - - - - - - -

Ethyle ne (C2H4) lb/hr - - - - - - - -

Acetylene (C2H2) lb/hr - - - - - - - -

Propa ne (C3H8) lb/hr - - - - - - - -

I-Buta ne (C4H10) lb/hr - - - - - - - -

N-Butane (C4H10) lb/hr - - - - - - - -

Penta ne + lb/hr - - - - - - - -

Benzene (C6H6) lb/hr - - - - - - - -

Tar (C10H8) lb/hr - - - - - - - -

Carbon (Solid) lb/hr - - - - - - - -

Sulfur (Solid) lb/hr - - - - - - - -

Olivine + MgO (Sol id) lb/hr2 495 495 - 495 --2

Methanol lb/hr - - - - - - - -

Ethanol lb/hr - - - - - - - -

Propa nol lb/hr - - - - - - - -

n-Butanol lb/hr - - - - - - - -

Penta nol+ lb/hr - - - - - - - -

Ash lb/hr- 1,690 1,690 - 1,690 ---

Char lb/hr 0 - - - - - -0

Wood lb/hr - - - - - - - -

Entha lpy Flow MMBTU/hr (1,191.0) (344.7) 0.4 (1.7) (2.2) (345.0) (535.6) (655.4)

97
Heat Stream No. MM BTU/hr Work Stream No. HP

QH305 1.58 WK305 1,729.7

QH306 2.64

QH307 5.91

QH308 5.33

QH309 12.75

QH311 9.08

COMPONENT UNITS 225F 328 329B 330A 330A-1 330A-2 385F 385F-1 430-1 514-MV

Total Flow l b/hr 9,080 108,286 253,679 253,679 253,679 253,679 36,261 36,261 59,646 3,298

Temp eratureF 1,596 1,179 90 161 316 600 578 1,139 1,170 140

Press ure Psia 31.4 19.0 14.7 20.0 20.0 19.0 33.7 28.7 19.0 43.0

Vapor Fra ction 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

Hydrogen l b/hr 114 1,423 ---- 1,302 1,302 8 0

Carbon Monoxide l b/hr 2,800 16,065 ---- 12,408 12,408 848 9

Nitrogen l b/hr0 257 187,845 187,845 187,845 187,845 242 242 15 0

Oxygen l b/hr-- 57,556 57,556 57,556 57,556 ----

Argon l b/hr-- 3,204 3,204 3,204 3,204 ----

Carbon Di oxide l b/hr 1,338 77,087 125 125 125 125 18,796 18,796 54,481 2,472

Water l b/hr 3,455 4,324 4,950 4,950 4,950 4,950 0 0 869 -

Hydrogen Sulfide (H2S) l b/hr2 11 ---- 0 0 1 8

SO2 l b/hr----------

Ammo nia (NH3) l b/hr 13 25 ---- 0 0 9 3

NO2 l b/hr----------

Methane (CH4) l b/hr 584 4,998 ---- 3,319 3,319 1,083 12

Ethane (C2H6) l b/hr 21 411 ---- 73 73 314 4

Ethylene (C2H4) l b/hr 290 645 ---- 99 99 250 6

Acetylene (C2H2) l b/hr 25 84 ------ 58 1

Propa ne (C3H8) l b/hr1 237 ---- 23 23 172 42

I-Buta ne (C4H10) l b/hr- 0 --------

N-Butane (C4H10) l b/hr 0 0 ---- 0 0 -0

Penta ne + l b/hr- 0 --------

Benzene (C6H6) l b/hr 24 24 ---- 0 0 -0

Tar (C10H8) l b/hr 71 71 ---- 0 0 --

Carbon (Solid) l b/hr----------

Sulfur (Solid) l b/hr----------

Olivine + MgO (Sol id) l b/hr 2 2 --------

Methanol l b/hr 335 2,121 ------ 1,054 732

Ethanol l b/hr3 470 ------ 464 3

Propa nol l b/hr- 22 ------ 22 -

n-Butanol l b/hr- 0 ---- 0 0 --

Penta nol+ l b/hr----------

Ash l b/hr 0 0 --------

Char l b/hr 0 0 --------

Wood l b/hr 0 0 --------

Entha lpy Flow MMBTU/h r (25.2) (324.7) (28.3) (23.9) (14.2) 3.9 (92.9) (83.9) (204.1) (11.5)

98
Heat Stream No. MM BTU/hr Work Stream No. HP

QH302 4.38

QH310 69.86

COMPONENT UNITS 373 385T 386 386A 393 513

Total Flow lb/hr 33,593 108,784 142,651 142,651 274 33,593

TemperatureF 564 578 575 1,700 590 72

Pressure Psia 35.0 33.7 33.7 28.7 59.7 40.0

Vapor Fraction 1.00 1.00 1.00 1.00 1.00 1.00

Hydrogen lb/hr 130 3,905 4,035 4,035 - 130

Carbon Monoxide lb/hr 2,289 37,224 39,513 39,513 - 2,289

Nitrogen lb/hr 32 726 758 758 - 32

Oxygen lb/hr------

Argon lb/hr------

Carbon Dioxide lb/hr 28,996 56,387 85,383 85,383 - 28,996

Water lb/hr 19 0 294 294 274 19

Hydrogen Sulfi de (H2S) lb/hr 25 1 26 26 - 25

SO2 lb/hr------

Ammonia (NH3) lb/hr3 0 4 4 -3

NO2 lb/hr------

Methane (CH4) lb/hr 1,100 9,956 11,056 11,056 - 1,100

Ethane (C2H6) lb/hr 83 219 302 302 - 83

Ethyl ene (C2H4) lb/hr 145 296 441 441 - 145

Acetyl ene (C2H2) lb/hr 7 -7 7 -7

Propane (C3H8) lb/hr 100 69 169 169 - 100

I-Butane (C4H10) lb/hr------

N-Butane (C4H10) lb/hr0 0 0 0 -0

Pentane + lb/hr------

Benzene (C6H6) lb/hr0 0 0 0 -0

Tar (C10H8) lb/hr0 0 0 0 -0

Carbon (Soli d) lb/hr------

Sulfur (Soli d) lb/hr------

Olivi ne + MgO (Solid) lb/hr------

Methanol lb/hr 487 - 487 487 - 487

Ethanol lb/hr 164 - 164 164 - 164

Propanol lb/hr 6 -6 6 -6

n-Butanol lb/hr0 0 0 0 -0

Pentanol + lb/hr------

Ash lb/hr------

Cha r lb/hr------

Wood lb/hr------

Entha lpy Fl ow MMBTU/hr (115.2) (278.8) (395.5) (325.7) (1.5) (119.6)

99
COMPONENT UNITS 225A 326 328 329 330 330A-2 332 386A

Total Flow l b/hr 236,707 12 108,286 12 379,359 253,679 361,965 142,651

TemperatureF 1,596 60 1,179 1,670 1,670 600 1,777 1,700

Pressure Psi a 31.4 29.0 19.0 25.7 25.7 19.0 18.0 28.7

Vapor Fracti on 1.00 0.00 1.00 0.00 1.00 1.00 1.00 1.00

Hydrogen l b/hr 2,963 - 1,423 - 17,862 - - 4,035

Carbon Monoxide l b/hr 73,002 - 16,065 - 198,960 - - 39,513

Nitrogen l b/hr 0 - 257 - 1,015 187,845 188,101 758

Oxygen l b/hr----- 57,556 7,680 -

Argon l b/hr----- 3,204 3,204 -

Carbon Dioxide l b/hr 34,886 - 77,087 - 90,677 125 124,583 85,383

Water l b/hr 90,078 - 4,324 - 64,279 4,950 38,305 294

Hydrogen Sul fide (H2S) l b/hr 54 - 11 - 80 -- 26

SO2 l b/hr------ 20 -

Ammonia (NH3) l b/hr 343 - 25 - 35 --4

NO2 l b/hr------ 69 -

Methane (CH4) l b/hr 15,232 - 4,998 - 5,258 - - 11,056

Ethane (C2H6) l b/hr 549 - 411 - 9 -- 302

Ethylene (C2H4) l b/hr 7,550 - 645 - 799 -- 441

Acetylene (C2H2) l b/hr 654 - 84 - 66 --7

Propane (C3H8) l b/hr 23 - 237 - 192 -- 169

I-Butane (C4H10) l b/hr-- 0 -----

N-Butane (C4H10) l b/hr 0 - 0 - 0 --0

Pentane + l b/hr-- 0 -----

Benzene (C6H6) l b/hr 617 - 24 - 6 --0

Tar (C10H8) l b/hr 1,852 - 71 - 19 --0

Carbon (Solid) l b/hr--------

Sul fur (Sol id) l b/hr--------

Oli vine + MgO (Soli d) l b/hr 48 12 2 12 48 - 2 -

Methanol l b/hr 8,734 - 2,121 - 0 -- 487

Ethanol l b/hr 73 - 470 - 0 -- 164

Propanol l b/hr-- 22 ----6

n-Butanol l b/hr-- 0 ----0

Pentanol + l b/hr--------

Ash l b/hr 0 - 0 - 0 ---

Char l b/hr 4 - 0 - 4 - 0 -

Wood l b/hr 0 - 0 - 0 ---

Enthalpy Flow MMBTU/hr (657.5) (0.0) (324.7) 0.0 (783.1) 3.9 (520.9) (325.7)

Heat Stream No. MM BTU/hr Work Stream No. HP

100
Heat Stream No. MM BTU/hr Work Stream No. HP

QCM301 9.92 WP302 10.9

COMPONENT UNITS 300 302B 308 313 415SCRUB

Tota l Flow l b/hr 379,359 29,562 341,407 351,310 1,514

Te mpera tu reF 140 130 130 137 136

Pressure Psi a 24.7 22.7 42.7 22.7 22.7

Va por Fraction 0.93 0.00 0.00 1.00 0.00

Hydrogen l b/hr 17,8620 0 17,862 -

Ca rbon Monoxide l b/hr 198,960 24 283 198,936 -

Nitrogen l b/hr 1,0150 0 1,015 -

Oxygen l b/hr-----

Argon l b/hr-----

Ca rbon Dioxide l b/hr 90,677 40 459 90,6380

Wa ter l b/hr 64,279 29,395 340,064 36,398 1,514

Hydrogen Sulfi de (H2S) l b/hr 800 1 80 -

SO2 l b/hr-----

Ammonia (NH3) l b/hr 35 15 172 200

NO2 l b/hr-----

Methane (CH4) l b/hr 5,258 7 85 5,250 -

Etha ne (C2H6) l b/hr900 8 -

Ethyl ene (C2H4) l b/hr 799 2 28 797 -

Acetyl ene (C2H2) l b/hr 660 2 66 -

Propane (C3H8) l b/hr 192 2 19 190 -

I-Butane (C4H10) l b/hr-----

N-Butane (C4H10) l b/hr000 0 -

Pe ntane + l b/hr-----

Benzene (C6H6) l b/hr61 15 5 -

Ta r (C10H8) l b/hr 19 18 209 0 -

Ca rbon (Sol id) l b/hr-----

Sulfur (Sol id) l b/hr-----

Olivi ne + MgO (Soli d) l b/hr 48 48 ---

Metha nol l b/hr000 0 -

Etha nol l b/hr000 0 -

Propanol l b/hr-----

n-Buta nol l b/hr-----

Pe ntanol + l b/hr-----

Ash l b/hr0 0 ---

Char l b/hr4 4 ---

Wood l b/hr0 0 ---

Enthalpy Flow MMBTU/hr (1,088.0 ) (199.3) (2,306.0) (897.8) (10.2)

101
COMPONENT UNITS 300 330 330B

Tota l Flow lb/hr 379,359 379,359 379,359

Tempera tureF 140 1,670 332

Pressure Psia 24.7 25.7 25.2

Va por Fra cti on 0.93 1.00 1.00

Hydrogen lb/hr 17,862 17,862 17,862

Ca rbon Monoxide lb/hr 198,960 198,960 198,960

Nitrogen lb/hr 1,015 1,015 1,015

Oxygen lb/hr---

Argon lb/hr---

Ca rbon Dioxide l b/hr 90,677 90,677 90,677

Wa ter lb/hr 64,279 64,279 64,279

Hydrogen Sul fide (H2S) l b/hr 80 80 80

SO2 lb/hr---

Ammonia (NH3) lb/hr 35 35 35

NO2 lb/hr---

Metha ne (CH4) lb/hr 5,258 5,258 5,258

Etha ne (C2H6) lb/hr9 9 9

Ethyl ene (C2H4) lb/hr 799 799 799

Acetylene (C2H2) lb/hr 66 66 66

Propa ne (C3H8) lb/hr 192 192 192

I-Butane (C4H10) lb/hr---

N-Butane (C4H10) lb/hr0 0 0

Penta ne + lb/hr---

Benzene (C6H6) lb/hr6 6 6

Ta r (C10H8) lb/hr 19 19 19

Ca rbon (Soli d) lb/hr---

Sulfur (Soli d) lb/hr---

Olivi ne + MgO (Solid) lb/hr 48 48 48

Metha nol lb/hr0 0 0

Etha nol lb/hr0 0 0

Propa nol lb/hr---

n-Buta nol lb/hr---

Penta nol + lb/hr---

Ash lb/hr0 0 0

Cha r lb/hr4 4 4

Wood lb/hr0 0 0

Entha lpy Fl ow MMBTU/hr (1,088.0) (783.1) (1,026.4)

Heat Stream No. MM BTU/hr Work Stream No. HP

QH303 7.22

QH304 5.66

102
COMPONENT UNITS 332 334

Tota l Fl ow lb/hr 361,965 361,965

Tempera tureF 1,777 549

Pressure Psia 18.0 17.0

Va por Fra cti on 1.00 1.00

Hydrogen lb/hr--

Ca rbon Monoxi de lb/hr--

Nitrogen lb/hr 188,101 188,101

Oxygen lb/hr 7,680 7,680

Argon lb/hr 3,204 3,204

Ca rbon Di oxide lb/hr 124,583 124,583

Wa ter lb/hr 38,305 38,305

Hydrogen Sulfide (H2S) lb/hr--

SO2 lb/hr 20 20

Ammonia (NH3) lb/hr--

NO2 lb/hr 69 69

Metha ne (CH4) lb/hr--

Etha ne (C2H6) lb/hr--

Ethylene (C2H4) lb/hr--

Acetylene (C2H2) lb/hr--

Propa ne (C3H8) lb/hr--

I-Butane (C4H10) lb/hr--

N-Butane (C4H10) lb/hr--

Penta ne + lb/hr--

Benzene (C6H6) lb/hr--

Ta r (C10H8) lb/hr--

Ca rbon (Solid) lb/hr--

Sulfur (Solid) lb/hr--

Olivi ne + MgO (Sol id) lb/hr 2 2

Metha nol lb/hr--

Etha nol lb/hr--

Propa nol lb/hr--

n-Buta nol lb/hr--

Penta nol+ lb/hr--

Ash lb/hr--

Char lb/hr 0 0

Wood lb/hr--

Entha lpy Flow MMBTU/hr (520.9) (655.4)

Heat Stream No. MM BTU/hr Work Stream No. HP

103
COMPONENT UNITS 313 410B-1 410C-1 410LRECY 415L 415V 466 478 483 483A

Total Flow l b/hr 351,310 320,378 318,986 35,120 35,980 350,451 315,330 709,768 709,768 709,768

TemperatureF 137 177 204 110 136 136 320 110 180 115

Press ure Psia 22.7 90.2 140.0 64.0 22.7 22.7 3,000.0 2,916.8 3,000.0 3,005.0

Vapor Fra ction 1.00 1.00 1.00 0.00 0.00 1.00 1.00 1.00 1.00 1.00

Hydrogen l b/hr 17,862 17,862 17,862 0 - 17,862 17,862 37,299 37,299 37,299

Carbon Monoxide l b/hr 198,936 198,936 198,936 0 - 198,936 198,936 311,136 311,136 311,136

Nitrogen l b/hr 1,015 1,015 1,015 0 - 1,015 1,015 6,119 6,119 6,119

Oxyge n l b/hr---- - -----

Argon l b/hr---- - -----

Carbon Di oxide l b/hr 90,638 90,638 90,638 2 0 90,639 90,637 282,118 282,118 282,118

Wate r l b/hr 36,398 5,463 4,072 35,111 35,978 35,531 420 1 1 1

Hydrogen Sulfide (H2S) l b/hr 80 80 80 0 - 80 80 8 8 8

SO2 l b/hr---- - -----

Ammonia (NH3) l b/hr 20 22 21 8 1 26 19 2 2 2

NO2 l b/hr---- - -----

Methane (CH4) l b/hr 5,250 5,250 5,250 0 - 5,250 5,250 70,307 70,307 70,307

Ethane (C2H6) l b/hr 8 8 8 0 - 8 8 1,097 1,097 1,097

Ethyle ne (C2H4) l b/hr 797 797 797 0 - 797 797 1,584 1,584 1,584

Acetylene (C2H2) l b/hr 66 66 66 0 - 66 66 - - -

Propa ne (C3H8) l b/hr 190 190 190 0 - 190 190 88 88 88

I-Buta ne (C4H10) l b/hr---- - -----

N-Butane (C4H10) l b/hr 0 0 0 0 - 0 0 0 0 0

Penta ne + l b/hr---- - -----

Benze ne (C6H6) l b/hr 5 5 5 0 - 5 5 0 0 0

Tar (C10H8) l b/hr 0 0 0 0 - 0 0 0 0 0

Carbon (Solid) l b/hr---- - -----

Sulfur (Solid) l b/hr---- - -----

Olivine + MgO (Sol id) l b/hr---- - -----

Methanol l b/hr 0 0 0 0 - 0 0 ---

Ethanol l b/hr 0 0 0 -- 0 0 ---

Propa nol l b/hr---- - -----

n-Butanol l b/hr---- - -- 0 0 0

Penta nol+ l b/hr---- - -----

Ash l b/hr---- - -----

Char l b/hr---- - -----

Wood l b/hr---- - -----

Entha lpy Flow MMBTU/hr (897.8) (714.6) (703.0) (240.8) (243.4) (892.9) (666.4) (1,757.7) (1,732.5) (1,756.0)

Heat Stream No. MM BTU/hr Work Stream No. HP

QAH410A 26.18 WK410A 15,296.7

QAH410B 6.37 WK410B 3,621.3

QAH410C 10.06 WK410C 5,078.0

QAH410D 6.53 WK410D 13,681.3

QAH410E 6.00 WK410E 13,640.1

QCH411A 7.94 WK410F 12,056.3

QCH411B 4.23 WK414 706.4

QCH411C 4.30

QCH411D 3.20

QCH411E 3.20

104
Heat Stream No. MM BTU/hr Work Stream No. HP

QAH413 26.16 WK412-1 (6,341.6)

QCH414 20.65 WK412-2 (9,036.2)

QH418 9.71

QH419 3.81

QH420 25.07

QH421 4.31

QH422 10.97

QH423 16.30

QR410 256.42

COMPONENT UNITS 385F 385T 426A 426B 426C 428 428-1 428B 471 472-GB 473A 474 477 477A 478 482-1 483 485 498 503 515 H2S-RECY

Total Flow l b/ hr 36,261 108,784 92,230 92,230 145,046 52,816 52,816 52,816 1,077,813 1,077,813 1,077,813 915,328 92,230 92,230 709,768 801,997 709,768 1,077,813 92,230 162,485 44,966 7,749

Te mpe rature F 578 578 595 950 578 30 316 950 611 175 110 110 95 316 110 110 180 595 950 110 164 80

Pressure Psi a 33.7 33.7 250.0 248.0 33.7 250.0 250.0 248.0 2,952.8 2,945.8 2,938.8 2,936.8 1,505.0 1,500.0 2,916.8 2,916.8 3,000.0 2,995.0 1,500.0 2,936.8 3,001.0 3,000.0

Va por Fraction 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.94 0.92 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00 0.00 0.00 0.00

Hydrogen l b/ hr 1,302 3,905 4,847 4,847 5,207 360 360 360 42,643 42,643 42,643 42,513 4,847 4,847 37,299 42,145 37,299 55,160 4,847 130 0 -

Ca rbon Monoxide l b/ hr 12,408 37,224 40,430 40,430 49,632 9,202 9,202 9,202 363,914 363,914 363,914 361,616 40,430 40,430 311,136 351,566 311,136 510,072 40,430 2,298 0 -

Nitrogen l b/ hr 242 726 795 795 968 173 173 173 7,134 7,134 7,134 7,102 795 795 6,119 6,914 6,119 7,134 795 33 0 -

Oxyge n l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Argon l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Ca rbon Dioxide l b/ hr 18,796 56,387 36,659 36,659 75,183 38,523 38,523 38,523 445,094 445,094 445,094 413,064 36,659 36,659 282,118 318,777 282,118 373,433 36,659 32,030 423 255

Water l b/ hr 0 0 0 0 0 - - - 7,261 7,261 7,261 50 0 0 1 1 1 421 0 7,211 - -

Hydrogen Sulfi de (H2S) l b/ hr 0 1 1 1 1 - - - 144 144 144 99 1 1 8 9 8 144 1 45 9 47

SO2 l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Ammonia (NH3) l b/ hr 0 0 0 0 0 - - - 27 27 27 12 0 0 2 2 2 27 0 15 6 0

NO2 l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Methane (CH4) l b/ hr 3,319 9,956 9,136 9,136 13,275 4,139 4,139 4,139 85,778 85,778 85,778 84,665 9,136 9,136 70,307 79,443 70,307 75,558 9,136 1,112 0 -

Ethane (C2H6) l b/ hr 73 219 142 142 292 149 149 149 1,789 1,789 1,789 1,702 142 142 1,097 1,239 1,097 1,105 142 87 0 -

Ethyl ene (C2H4) l b/ hr 99 296 206 206 395 189 189 189 2,382 2,382 2,382 2,229 206 206 1,584 1,790 1,584 2,382 206 153 1 -

Ace tyl ene (C2H2) l b/ hr - - - - - - - - 66 66 66 58 - - - - - 66 - 8 0 -

Propa ne (C3H8) l b/ hr 23 69 11 11 92 81 81 81 631 631 631 352 11 11 88 100 88 376 11 279 98 -

I-Butane (C4H10) l b/ hr - - - - - - - - - - - - - - - - - - - - - -

N-Buta ne (C4H10) l b/ hr 0 0 0 0 0 - - - 2 2 2 1 0 0 0 1 0 2 0 1 1 -

Pe nta ne + l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Benzene (C6H6) l b/ hr 0 0 0 0 0 - - - 5 5 5 0 0 0 0 0 0 5 0 5 0 -

Ta r (C10H8) l b/ hr 0 0 0 0 0 - - - 0 0 0 0 0 0 0 0 0 0 0 0 - -

Ca rbon (Sol id) l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Sulfur (Sol id) l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Olivi ne + MgO (Soli d) l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Methanol l b/ hr - - - - - - - - 62,811 62,811 62,811 1,301 - - - - - 51,214 - 61,510 43,836 7,377

Ethanol l b/ hr - - - - - - - - 52,192 52,192 52,192 533 - - - - - 439 - 51,659 370 69

Propa nol l b/ hr - - - - - - - - 5,340 5,340 5,340 22 - - - - - 0 - 5,318 - 0

n-Butanol l b/ hr 0 0 0 0 0 - - - 7 7 7 0 0 0 0 0 0 0 0 7 - -

Pe nta nol + l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Ash l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Char l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Wood l b/ hr - - - - - - - - - - - - - - - - - - - - - -

Entha lpy Flow MMBTU/hr (92.9) (278.8) (206.4) (190.2) (371.8) (173.2) (168.9) (157.9) (2,601.1) (2,859.7) (2,891.1) (2,366.2) (228.4) (218.7) (1,757.7) (1,986.2) (1,732.5) (2,344.7) (189.8) (524.9) (142.3) (25.1)

105
Heat Stream No. MM BTU/hr Work Stream No. HP

QCM401 79.49

COMPONENT UNITS 428 474 482-1 CO2-H2S

Total Fl ow lb/hr 52,816 915,328 801,997 60,514

TemperatureF 30 110 110 110

Press ure Psia 250.0 2,936.8 2,916.8 2,916.8

Vapor Fra ction 1.00 1.00 1.00 1.00

Hydrogen lb/hr 360 42,513 42,145 8

Carbon Monoxide lb/hr 9,202 361,616 351,566 848

Nitrogen lb/hr 173 7,102 6,914 15

Oxyge n lb/hr - - - -

Argon lb/hr - - - -

Carbon Di oxide lb/hr 38,523 413,064 318,777 55,764

Wate r lb/hr- 50 1 49

Hydrogen Sulfide (H2S) lb/hr- 99 9 89

SO2 lb/hr - - - -

Ammonia (NH3) lb/hr- 12 2 9

NO2 lb/hr - - - -

Methane (CH4) lb/hr 4,139 84,665 79,443 1,083

Ethane (C2H6) lb/hr 149 1,702 1,239 314

Ethyle ne (C2H4) lb/hr 189 2,229 1,790 250

Acetylene (C2H2) lb/hr- 58 - 58

Propa ne (C3H8) lb/hr 81 352 100 172

I-Buta ne (C4H10) lb/hr - - - -

N-Butane (C4H10) lb/hr- 1 1 -

Penta ne + lb/hr - - - -

Benze ne (C6H6) lb/hr- 0 0 -

Tar (C10H8) lb/hr- 0 0 -

Carbon (Solid) lb/hr - - - -

Sulfur (Solid) lb/hr - - - -

Olivine + MgO (Sol id) lb/hr - - - -

Methanol lb/hr- 1,301 - 1,301

Ethanol lb/hr- 533 - 533

Propa nol lb/hr- 22 - 22

n-Butanol lb/hr- 0 0 -

Penta nol+ lb/hr - - - -

Ash lb/hr - - - -

Char lb/hr - - - -

Wood lb/hr - - - -

Entha lpy Flow MMBTU/hr (1.7) (23.7) (19.9) (2.3)

106
COMPONENT UNITS 430 CO2-H2S H2O-MKUP H2S-RCH2

Total Flow lb/hr 59,646 60,514 854 1,722

TemperatureF 94 110 90 110

Pressure Psia 23.0 2,916.8 30.0 30.0

Vapor Fraction- 0.99 1.00 - 0.87

Hydrogen lb/hr8 8 --

Carbon Monoxide lb/hr 848 848 --

Nitrogen lb/hr 15 15 --

Oxygen lb/hr----

Argon lb/hr----

Carbon Dioxide lb/hr 54,481 55,764 - 1,283

Water lb/hr 869 49 854 34

Hydrogen Sulfi de (H2S) lb/hr1 89 - 88

SO2 lb/hr----

Ammonia (NH3) lb/hr9 9 -0

NO2 lb/hr----

Methane (CH4) lb/hr 1,083 1,083 --

Ethane (C2H6) lb/hr 314 314 --

Ethyl ene (C2H4) lb/hr 250 250 --

Acetyl ene (C2H2) lb/hr 58 58 --

Propane (C3H8) lb/hr 172 172 --

I-Butane (C4H10) lb/hr----

N-Butane (C4H10) lb/hr----

Pentane + lb/hr----

Benzene (C6H6) lb/hr----

Tar (C10H8) lb/hr----

Carbon (Soli d) lb/hr----

Sulfur (Soli d) lb/hr----

Olivi ne + MgO (Solid) lb/hr----

Methanol lb/hr 1,054 1,301 - 247

Ethanol lb/hr 464 533 - 69

Propanol lb/hr 22 22 -0

n-Butanol lb/hr----

Pentanol + lb/hr----

Ash lb/hr----

Cha r lb/hr----

Wood lb/hr----

Entha lpy Fl ow MMBTU/hr (2.2) (2.3) (0.1) (0.1)

Heat Stream No. MM BTU/hr Work Stream No. HP

107
COMPONENT UNITS CO2-H2S1 H2S-RCH2 H2S-RECY COLDMEOH

Total Fl ow lb/hr 1,103 1,722 7,749 7,130

TemperatureF 80 110 80 140

Press ure Psia 28.0 30.0 3,000.0 43.0

Vapor Fracti on- 0.94 0.87 --

Hydrogen lb/hr----

Carbon Monoxide lb/hr----

Nitrogen lb/hr----

Oxygen lb/hr----

Argon lb/hr----

Carbon Di oxide lb/hr 1,028 1,283 255 -

Water lb/hr 34 34 --

Hydrogen Sul fi de (H2S) lb/hr 41 88 47 -

SO2 lb/hr----

Ammonia (NH3) lb/hr- 0 0 -

NO2 lb/hr----

Methane (CH4) lb/hr----

Ethane (C2H6) lb/hr----

Ethylene (C2H4) lb/hr----

Acetylene (C2H2) lb/hr----

Propane (C3H8) lb/hr----

I-Butane (C4H10) lb/hr----

N-Butane (C4H10) lb/hr----

Pentane + lb/hr----

Benzene (C6H6) lb/hr----

Tar (C10H8) lb/hr----

Carbon (Soli d) lb/hr----

Sulfur (Soli d) lb/hr----

Oli vine + MgO (Soli d) lb/hr----

Methanol lb/hr- 247 7,377 7,130

Ethanol lb/hr- 69 69 -

Propanol lb/hr- 0 0 -

n-Butanol lb/hr----

Pentanol+ lb/hr----

Ash lb/hr----

Char lb/hr----

Wood lb/hr----

Enthal py Fl ow MMBTU/hr (0.0) (0.1) (0.3) (0.2)

Heat Stream No. MM BTU/hr Work Stream No. HP

108
COMPONENT UNITS CO2-H2S1 CO2-H2S2 CO2-VENT 422 424

Tota l Flow lb/hr 1,103 1,103 1,062 94 39

Tempera tureF 80 150 149 90 150

Pressure Psi a 28.0 23.0 18.0 14.7 14.7

Va por Fra cti on- 0.94 1.00 1.00 1.00 -

Hydrogen lb/hr-----

Ca rbon Monoxide lb/hr-----

Nitrogen lb/hr--- 70 -

Oxygen lb/hr--- 21 -

Argon lb/hr--- 1 -

Ca rbon Dioxide lb/hr 1,028 1,028 1,028 0 -

Wa ter lb/hr 34 34 34 2 -

Hydrogen Sul fide (H2S) l b/hr 41 41 0 --

SO2 lb/hr-----

Ammonia (NH3) lb/hr-----

NO2 lb/hr-----

Metha ne (CH4) lb/hr-----

Etha ne (C2H6) lb/hr-----

Ethyl ene (C2H4) lb/hr-----

Acetylene (C2H2) lb/hr-----

Propa ne (C3H8) lb/hr-----

I-Butane (C4H10) lb/hr-----

N-Butane (C4H10) lb/hr-----

Penta ne + lb/hr-----

Benzene (C6H6) lb/hr-----

Ta r (C10H8) lb/hr-----

Ca rbon (Soli d) lb/hr-----

Sulfur (Soli d) lb/hr---- 39

Olivi ne + MgO (Solid) lb/hr-----

Metha nol lb/hr-----

Etha nol lb/hr-----

Propa nol lb/hr-----

n-Buta nol lb/hr-----

Penta nol + lb/hr-----

Ash lb/hr-----

Cha r lb/hr-----

Wood lb/hr-----

Entha lpy Fl ow MMBTU/hr (0.0) (0.0) (0.0) (0.0) 0.0

Heat Stream No. MM BTU/hr Work Stream No. HP

QH409 0.05

QCM405 0.01

109
Heat Stream No. MM BTU/hr Work Stream No. HP

QH401 127.70

QH402 73.55

QH415 19.56

QH416 1.01

QH417 5.35

QH424 2.37

QH425 4.52

QH426 10.14

QH427 20.42

COMPONENT UNITS 430 430A 430-1 470 470A 471 471B 472 472-GB 485

Total Flow l b/hr 59,646 59,646 59,646 1,077,813 1,077,813 1,077,813 1,077,813 1,077,813 1,077,813 1,077,813

TemperatureF 94 316 1,170 182 316 611 332 200 175 595

Press ure Psia 23.0 23.0 19.0 3,000.0 2,995.0 2,952.8 2,947.8 2,945.8 2,945.8 2,995.0

Vapor Fra ction 0.99 1.00 1.00 1.00 1.00 1.00 1.00 0.93 0.94 1.00

Hydrogen l b/hr 8 8 8 55,160 55,160 42,643 42,643 42,643 42,643 55,160

Carbon Monoxide l b/hr 848 848 848 510,072 510,072 363,914 363,914 363,914 363,914 510,072

Nitrogen l b/hr 15 15 15 7,134 7,134 7,134 7,134 7,134 7,134 7,134

Oxyge n l b/hr---- - -----

Argon l b/hr---- - -----

Carbon Di oxide l b/hr 54,481 54,481 54,481 373,433 373,433 445,094 445,094 445,094 445,094 373,433

Water l b/hr 869 869 869 421 421 7,261 7,261 7,261 7,261 421

Hydrogen Sulfide (H2S) l b/hr 1 1 1 144 144 144 144 144 144 144

SO2 l b/hr---- - -----

Ammonia (NH3) l b/hr 9 9 9 27 27 27 27 27 27 27

NO2 l b/hr---- - -----

Methane (CH4) l b/hr 1,083 1,083 1,083 75,558 75,558 85,778 85,778 85,778 85,778 75,558

Ethane (C2H6) l b/hr 314 314 314 1,105 1,105 1,789 1,789 1,789 1,789 1,105

Ethyle ne (C2H4) l b/hr 250 250 250 2,382 2,382 2,382 2,382 2,382 2,382 2,382

Acetylene (C2H2) l b/hr 58 58 58 66 66 66 66 66 66 66

Propa ne (C3H8) l b/hr 172 172 172 376 376 631 631 631 631 376

I-Buta ne (C4H10) l b/hr---- - -----

N-Butane (C4H10) l b/hr--- 2 2 2 2 2 2 2

Penta ne + l b/hr---- - -----

Benzene (C6H6) l b/hr--- 5 5 5 5 5 5 5

Tar (C10H8) l b/hr--- 0 0 0 0 0 0 0

Carbon (Solid) l b/hr---- - -----

Sulfur (Solid) l b/hr---- - -----

Olivine + MgO (Sol id) l b/hr---- - -----

Methanol l b/hr 1,054 1,054 1,054 51,214 51,214 62,811 62,811 62,811 62,811 51,214

Ethanol l b/hr 464 464 464 439 439 52,192 52,192 52,192 52,192 439

Propa nol l b/hr 22 22 22 0 0 5,340 5,340 5,340 5,340 0

n-Butanol l b/hr--- 0 0 7 7 7 7 0

Penta nol+ l b/hr---- - -----

Ash l b/hr---- - -----

Char l b/hr---- - -----

Wood l b/hr---- - -----

Entha lpy Flow MMBTU/hr (2.2) (2.2) (2.0) (25.7) (24.9) (26.0) (27.3) (28.4) (28.6) (23.4)

110
Heat Stream No. MM BTU/hr Work Stream No. HP

QH403 25.04

QH404 15.30

QH405 15.53

QH406 15.96

QH407 11.10

QH408 11.39

COMPONENT UNITS 410A-1 410A-2 410D-1 410D-2 410E-1 410E-2

Total Flow lb/hr 350,451 350,451 317,529 317,529 315,903 315,903

TemperatureF 388 170 360 170 359 170

Press ure Psia 70.0 68.0 425.0 423.0 1,215.0 1,213.0

Vapor Fra ction 1.00 0.98 1.00 1.00 1.00 1.00

Hydrogen lb/hr 17,862 17,862 17,862 17,862 17,862 17,862

Carbon Monoxide lb/hr 198,936 198,936 198,936 198,936 198,936 198,936

Nitrogen lb/hr 1,015 1,015 1,015 1,015 1,015 1,015

Oxygen lb/hr- - - - - -

Argon lb/hr- - - - - -

Carbon Di oxide lb/hr 90,639 90,639 90,638 90,638 90,637 90,637

Water lb/hr 35,531 35,531 2,616 2,616 991 991

Hydrogen Sulfide (H2S) lb/hr 80 80 80 80 80 80

SO2 lb/hr- - - - - -

Ammoni a (NH3) lb/hr 26 26 21 21 20 20

NO2 lb/hr- - - - - -

Methane (CH4) lb/hr 5,250 5,250 5,250 5,250 5,250 5,250

Ethane (C2H6) lb/hr 8 8 8 8 8 8

Ethylene (C2H4) lb/hr 797 797 797 797 797 797

Acetylene (C2H2) lb/hr 66 66 66 66 66 66

Propa ne (C3H8) lb/hr 190 190 190 190 190 190

I-Buta ne (C4H10) lb/hr- - - - - -

N-Butane (C4H10) lb/hr 0 0 0 0 0 0

Penta ne + lb/hr- - - - - -

Benzene (C6H6) lb/hr 5 5 5 5 5 5

Tar (C10H8) lb/hr 0 0 0 0 0 0

Carbon (Solid) lb/hr- - - - - -

Sulfur (Solid) lb/hr- - - - - -

Olivine + MgO (Sol id) lb/hr- - - - - -

Methanol lb/hr 0 0 0 0 0 0

Ethanol lb/hr 0 0 0 0 0 0

Propa nol lb/hr- - - - - -

n-Butanol lb/hr- - - - - -

Penta nol+ lb/hr- - - - - -

Ash lb/hr- - - - - -

Char lb/hr- - - - - -

Wood lb/hr- - - - - -

Entha lpy Flow MMBTU/hr (8.5) (9.0) (6.7) (7.0) (6.6) (6.9)

111
COMPONENT UNITS 503 505 506 507 513 513L 513V 531 FLUSH

Total Flow lb/hr 162,485 129,421 129,421 122,834 33,593 15,867 529 33,064 9,808

TemperatureF 110 67 258 229 72 140 140 67 202

Pressure Psi a 2,936.8 60.0 55.0 55.0 40.0 40.0 40.0 60.0 45.0

Vapor Fraction--- 1.00 1.00 1.00 - 1.00 1.00 1.00

Hydrogen lb/hr 13000 0 1300 0 1300

Carbon Monoxi de lb/hr 2,298 11 11 11 2,2890 2 2,2872

Nitrogen lb/hr 3300 0 320 0 320

Oxygen lb/hr---------

Argon lb/hr---------

Carbon Dioxi de lb/hr 32,030 3,487 3,487 3,487 28,996 69 454 28,543 523

Water lb/hr 7,211 7,202 7,202 614 19 6,578 9 100

Hydrogen Sulfide (H2S) l b/hr 45 21 21 21 252 1 243

SO2 lb/hr---------

Ammoni a (NH3) lb/hr 15 12 12 1232032

NO2 lb/hr---------

Methane (CH4) lb/hr 1,112 15 15 15 1,1000 2 1,0972

Ethane (C2H6) lb/hr 8755 5 830 1 821

Ethyl ene (C2H4) lb/hr 15399 9 1450 1 1441

Acetyl ene (C2H2) lb/hr82227 0 070

Propane (C3H8) lb/hr 279 182 182 182 100 24 4 96 27

I-Butane (C4H10) lb/hr---------

N-Butane (C4H10) lb/hr11110 0 000

Pentane + lb/hr---------

Benzene (C6H6) lb/hr55550 0 000

Tar (C10H8) lb/hr0000 0 -- 0 -

Carbon (Sol id) lb/hr---------

Sulfur (Sol id) lb/hr---------

Olivine + MgO (Sol id) lb/hr---------

Methanol lb/hr 61,510 61,078 61,078 61,078 487 9,069 55 432 9,123

Ethanol lb/hr 51,659 51,495 51,495 51,495 164 76 0 164 76

Propanol lb/hr 5,318 5,311 5,311 5,311 6 --60

n-Butanol lb/hr7777 0 -- 0 -

Pentanol+ lb/hr---------

Ash lb/hr---------

Char lb/hr---------

Wood lb/hr---------

Enthalpy Flow MMBTU/hr (5.2) (4.1) (3.4) (3.0) (1.2) (0.7) (0.0) (1.2) (0.3)

Heat Stream No. MM BTU/hr Work Stream No. HP

QAS513 11.71

112
Heat Stream No. MM BTU/hr Work Stream No. HP

QH501 4.81 WP514 372.4

QH502 0.76

QAH504 35.16

QAH505 94.49

QCH591 0.05

QCH593 0.40

QRH504 (32.40)

QRH505 (77.56)

QAH514 25.88

COMPONENT UNITS 507 510 511 512 514-M 514-M0 514-MV 515 534 536 590B 590C 592A 592B 592C COLDMEOH FLUSH

Total Flow l b/hr 122,834 116,386 6,448 65,385 55,578 55,578 3,298 44,966 102,663 102,663 6,448 6,448 51,000 51,000 51,000 7,314 9,808

Te mpera tureF 229 214 265 202 202 140 140 164 654 287 120 110 230 120 110 140 202

Pressure Psi a 55.0 45.0 45.0 45.0 45.0 43.0 43.0 3,001.0 64.7 54.7 40.0 35.0 45.0 40.0 35.0 43.0 45.0

Va por Fraction 1.00 1.00 0.00 1.00 1.00 0.05 1.00 1.00 1.00

Hydroge n l b/hr0 0 -000 0 0 -------00

Ca rbon Monoxide l b/hr 11 11 - 11 10 109 0 -------02

Nitrogen l b/hr0 0 -000 0 0 -------00

Oxygen l b/hr-----------------

Argon l b/hr-----------------

Ca rbon Dioxide l b/hr 3,487 3,487 - 3,487 2,964 2,964 2,472 423 ------- 69 523

Water l b/hr 614 0 61400 0 - - 102,663 102,663 614 61400 0 -0

Hydroge n Sulfi de (H2S) l b/hr 21 21 0 21 18 188 9 --0 0 ---13

SO2 l b/hr-----------------

Ammonia (NH3) l b/hr 12 12 0 12 10 103 6 --0 0 ---12

NO2 l b/hr-----------------

Methane (CH4) l b/hr 15 15 - 15 13 13 12 0 -------02

Ethane (C2H6) l b/hr5 5 -544 4 0 -------01

Ethyl ene (C2H4) l b/hr990988 6 1 --0 0 ---01

Ace tyl ene (C2H2) l b/hr2 2 -211 1 0 -------00

Propane (C3H8) l b/hr 182 182 0 182 155 155 42 98 --0 0 --- 16 27

I-Butane (C4H10) l b/hr-----------------

N-Butane (C4H10) l b/hr110111 0 1 --0 0 00000

Pe nta n e + l b/hr-----------------

Benzene (C6H6) l b/hr532000 0 0 --2 2 33300

Ta r (C10H8) l b/hr 0 - 0 -------0 0 -----

Ca rbon (Sol id) l b/hr-----------------

Sulfur (Sol id) l b/hr-----------------

Olivi ne + MgO (Soli d) l b/hr-----------------

Methanol l b/hr 61,078 61,077 0 60,822 51,699 51,699 732 43,836 --0 0 256 256 256 7,130 9,123

Ethanol l b/hr 51,495 50,980 515 510 433 4333 370 -- 515 515 50,470 50,470 50,470 60 76

Propa n ol l b/hr 5,311 53 5,25800 0 --- - 5,258 5,258 53 53 53 -0

n-Butanol l b/hr70 7 -------7 7 -----

Pe nta n ol + l b/hr-----------------

Ash l b/hr-----------------

Ch a r l b/hr-----------------

Wood l b/hr-----------------

Enthalpy Flow MMBTU/hr (3.0) (2.9) (0.2) (1.8) (1.5) (1.8) (0.1) (1.4) (5.7) (6.8) (0.2) (0.2) (1.3) (1.3) (1.3) (0.2) (0.3)

113
Heat Stream No. MM BTU/hr Work Stream No. HP

QH503 12.14

QH506 10.42

QH507 1.03

QH508 0.88

QH509 0.48

QH510 1.52

COMPONENT UNITS 505 505B 506 513L 513L-G1 513-LG2A

Total Flow lb/hr 129,421 129,421 129,421 15,867 15,867 15,867

TemperatureF 67 238 258 140 590 316

Press ure Psia 60.0 55.0 55.0 40.0 35.0 35.0

Vapor Fra ction 1.00 1.00 1.00 1.00

Hydrogen lb/hr 0 0 0 0 0 0

Carbon Monoxide lb/hr 11 11 11 0 0 0

Nitrogen lb/hr 0 0 0 0 0 0

Oxyge n lb/hr - - - - - -

Argon lb/hr - - - - - -

Carbon Di oxid e lb/hr 3,487 3,487 3,487 69 69 69

Wate r lb/hr 7,202 7,202 7,202 6,578 6,578 6,578

Hydrogen Sulfide (H2S) lb/hr 21 21 21 2 2 2

SO2 lb/hr - - - - - -

Ammonia (NH3) lb/hr 12 12 12 2 2 2

NO2 lb/hr - - - - - -

Methane (CH4) lb/hr 15 15 15 0 0 0

Ethane (C2H6) lb/hr 5 5 5 0 0 0

Ethyle ne (C2H4) lb/hr 9 9 9 0 0 0

Acetylene (C2H2) lb/hr 2 2 2 0 0 0

Propa ne (C3H8) lb/hr 182 182 182 24 24 24

I-Buta ne (C4H10) lb/hr - - - - - -

N-Butane (C4H10) lb/hr 1 1 1 0 0 0

Penta ne + lb/hr - - - - - -

Benze ne (C6H6) lb/hr 5 5 5 0 0 0

Tar (C10H8) lb/hr 0 0 0 - - -

Carbon (Solid) lb/hr - - - - - -

Sulfur (Solid) lb/hr - - - - - -

Olivine + MgO (Sol id) lb/hr - - - - - -

Methanol lb/hr 61,078 61,078 61,078 9,069 9,069 9,069

Ethanol lb/hr 51,495 51,495 51,495 76 76 76

Propa nol lb/hr 5,311 5,311 5,311 ---

n-Butanol lb/hr 7 7 7 - - -

Penta nol+ lb/hr - - - - - -

Ash lb/hr - - - - - -

Char lb/hr - - - - - -

Wood lb/hr - - - - - -

Entha lpy Flow MMBTU/hr (4.1) (3.4) (3.4) (0.7) (0.6) (0.6)

114
Heat Stream No. MM BTU/hr Work Stream No. HP

QH605 6.50 WP601 13.8

QH606 6.50 WP603 18.7

QH607 9.98 WP604 1,046.7

COMPONENT UNITS 218 536 616 618 618A 620 620B 618ROBRN 628 631 638 640 BRINEWWT ROBRNRCY RODIRTY SELSTRET TOROBRIN

Total Flow l b/hr 243 102,663 245,642 101,798 101,798 76,349 76,349 25,450 476,788 476,788 476,791 476,791 7,710 23,858 6,361 52,134 31,811

TemperatureF 66 287 136 60 60 60 239 60 198 232 232 235 66 66 90 289 66

Press ure Psia 75.0 54.7 24.7 14.7 75.0 29.7 24.7 15.0 39.7 21.7 21.7 1,351.0 75.0 75.0 14.7 54.7 14.7

Vapor Fra ction

Hydrogen l b/hr- - -- - -- - - - - - --- - -

Carbon Monoxide l b/hr- - -- - -- - - - - - --- - -

Nitrogen l b/hr- - -- - -- - - - - - --- - -

Oxyge n l b/hr- - -- - -- - - - - - --- - -

Argon l b/hr- - -- - -- - - - - - --- - -

Carbon Di oxide l b/hr- - -- - -- - - - - - --- - -

Wate r l b/hr 243 102,663 245,642 101,798 101,798 76,349 76,349 25,450 476,788 476,788 476,791 476,791 7,710 23,858 6,361 52,134 31,811

Hydrogen Sulfi de (H2S) l b/hr- - -- - -- - - - - - --- - -

SO2 l b/hr- - -- - -- - - - - - --- - -

Ammonia (NH3) l b/hr- - -- - -- - - - - - --- - -

NO2 l b/hr- - -- - -- - - - - - --- - -

Methane (CH4) l b/hr- - -- - -- - - - - - --- - -

Ethane (C2H6) l b/hr- - -- - -- - - - - - --- - -

Ethyle ne (C2H4) l b/hr- - -- - -- - - - - - --- - -

Acetylene (C2H2) l b/hr- - -- - -- - - - - - --- - -

Propa ne (C3H8) l b/hr- - -- - -- - - - - - --- - -

I-Buta ne (C4H10) l b/hr- - -- - -- - - - - - --- - -

N-Butane (C4H10) l b/hr- - -- - -- - - - - - --- - -

Penta ne + l b/hr- - -- - -- - - - - - --- - -

Benze ne (C6H6) l b/hr- - -- - -- - - - - - --- - -

Tar (C10H8) l b/hr- - -- - -- - - - - - --- - -

Carbon (Solid) l b/hr- - -- - -- - - - - - --- - -

Sulfur (Solid) l b/hr- - -- - -- - - - - - --- - -

Olivine + MgO (Sol id) l b/hr- - -- - -- - - - - - --- - -

Methanol l b/hr- - -- - -- - - - - - --- - -

Ethanol l b/hr- - -- - -- - - - - - --- - -

Propa nol l b/hr- - -- - -- - - - - - --- - -

n-Butanol l b/hr- - -- - -- - - - - - --- - -

Penta nol+ l b/hr- - -- - -- - - - - - --- - -

Ash l b/hr- - -- - -- - - - - - --- - -

Char l b/hr- - -- - -- - - - - - --- - -

Wood l b/hr- - -- - -- - - - - - --- - -

Entha lpy Flow MMBTU/hr (0.0) (6.8) (16.6) (7.0) (7.0) (5.2) (5.1) (1.7) (32.0) (31.8) (31.8) (31.8) (0.5) (1.6) (0.4) (3.5)

115

(2.2)
Heat Stream No. MM BTU/hr Work Stream No. HP

QAH601 252.27 WTURB1 (23,244.0)

QH602 3.92 WTURB2 (47,755.1)

QH604 4.38 WP602 9.0

QH613 74.62

COMPONENT UNITS 200 393 534 602C 607 608 614A 615 616 SELEXSTM

Total Flow l b/hr 66,542 274 102,663 467,255 467,255 467,255 245,642 245,642 245,642 52,134

TemperatureF 649 654 654 1,000 1,000 698 234 136 136 654

Press ure Psia 25.0 64.7 64.7 345.0 1,321.0 350.0 3.5 2.6 24.7 64.7

Vapor Fra ction 1.00 1.00 1.00 1.00 1.00 1.00 1.00 1.00

Hydrogen l b/hr---- - -----

Carbon Monoxide l b/hr---- - -----

Nitrogen l b/hr---- - -----

Oxyge n l b/hr---- - -----

Argon l b/hr---- - -----

Carbon Di oxide l b/hr---- - -----

Wate r l b/hr 66,542 274 102,663 467,255 467,255 467,255 245,642 245,642 245,642 52,134

Hydrogen Sulfide (H2S) l b/hr---- - -----

SO2 l b/hr---- - -----

Ammonia (NH3) l b/hr---- - -----

NO2 l b/hr---- - -----

Methane (CH4) l b/hr---- - -----

Ethane (C2H6) l b/hr---- - -----

Ethyle ne (C2H4) l b/hr---- - -----

Acetylene (C2H2) l b/hr---- - -----

Propa ne (C3H8) l b/hr---- - -----

I-Buta ne (C4H10) l b/hr---- - -----

N-Butane (C4H10) l b/hr---- - -----

Penta ne + l b/hr---- - -----

Benze ne (C6H6) l b/hr---- - -----

Tar (C10H8) l b/hr---- - -----

Carbon (Solid) l b/hr---- - -----

Sulfur (Solid) l b/hr---- - -----

Olivine + MgO (Sol id) l b/hr---- - -----

Methanol l b/hr---- - -----

Ethanol l b/hr---- - -----

Propa nol l b/hr---- - -----

n-Butanol l b/hr---- - -----

Penta nol+ l b/hr---- - -----

Ash l b/hr---- - -----

Char l b/hr---- - -----

Wood l b/hr---- - -----

Entha lpy Flow MMBTU/hr (3.7) (0.0) (5.7) (25.0) (25.1) (25.7) (14.0) (16.6) (16.6) (2.9)

116
Heat Stream No. MM BTU/hr Work Stream No. HP

QCH603 0.47

QH612 147.91

COMPONENT UNITS 600 603 604 606 607 643 643M2

Total Flow lb/hr 9,536 9,536 9,536 467,255 467,255 476,791 30,758

TemperatureF 113 160 581 580 1,000 581 581

Press ure Psia 14.7 1,331.0 1,336.0 1,326.0 1,321.0 1,336.0 1,336.0

Vapor Fra ction 1.00 1.00 0.98 0.98

Hydrogen lb/hr - - - - - - -

Carbon Monoxide lb/hr - - - - - - -

Nitrogen lb/hr - - - - - - -

Oxyge n lb/hr - - - - - - -

Argon lb/hr - - - - - - -

Carbon Di oxid e lb/hr - - - - - - -

Wate r lb/hr 9,536 9,536 9,536 467,255 467,255 476,791 30,758

Hydrogen Sulfide (H2S) lb/hr-------

SO2 lb/hr - - - - - - -

Ammonia (NH3) lb/hr - - - - - - -

NO2 lb/hr - - - - - - -

Methane (CH4) lb/hr - - - - - - -

Ethane (C2H6) lb/hr - - - - - - -

Ethyle ne (C2H4) lb/hr - - - - - - -

Acetylene (C2H2) lb/hr - - - - - - -

Propa ne (C3H8) lb/hr - - - - - - -

I-Buta ne (C4H10) lb/hr - - - - - - -

N-Butane (C4H10) lb/hr - - - - - - -

Penta ne + lb/hr - - - - - - -

Benze ne (C6H6) lb/hr - - - - - - -

Tar (C10H8) lb/hr - - - - - - -

Carbon (Solid) lb/hr - - - - - - -

Sulfur (Solid) lb/hr - - - - - - -

Olivine + MgO (Sol id) lb/hr - - - - - - -

Methanol lb/hr - - - - - - -

Ethanol lb/hr - - - - - - -

Propa nol lb/hr - - - - - - -

n-Butanol lb/hr - - - - - - -

Penta nol+ lb/hr - - - - - - -

Ash lb/hr - - - - - - -

Char lb/hr - - - - - - -

Wood lb/hr - - - - - - -

Entha lpy Flow MMBTU/hr (0.7) (0.6) (0.6) (26.6) (25.1) (27.2) (1.8)

117
Heat Stream No. MM BTU/hr Work Stream No. HP

QH608 63.73

QH609 26.06

QH610 38.11

QH611 17.67

COMPONENT UNITS 640 641 642M2 643M2

Total Flow l b/hr 476,791 476,791 30,758 30,758

TemperatureF 235 582 581 581

Press ure Psia 1,351.0 1,346.0 1,336.0 1,336.0

Vapor Fra ction 0.00 0.98

Hydrogen l b/hr----

Carbon Monoxide l b/hr----

Nitrogen l b/hr----

Oxyge n l b/hr----

Argon l b/hr----

Carbon Di oxide l b/hr----

Water l b/hr 476,791 476,791 30,758 30,758

Hydrogen Sulfide (H2S) l b/hr----

SO2 l b/hr----

Ammonia (NH3) l b/hr----

NO2 l b/hr----

Methane (CH4) l b/hr----

Ethane (C2H6) l b/hr----

Ethyle ne (C2H4) l b/hr----

Acetylene (C2H2) l b/hr----

Propa ne (C3H8) l b/hr----

I-Buta ne (C4H10) l b/hr----

N-Butane (C4H10) l b/hr----

Penta ne + l b/hr----

Benze ne (C6H6) l b/hr----

Tar (C10H8) l b/hr----

Carbon (Solid) l b/hr----

Sulfur (Solid) l b/hr----

Olivine + MgO (Sol id) l b/hr----

Methanol l b/hr----

Ethanol l b/hr----

Propa nol l b/hr----

n-Butanol l b/hr----

Penta nol+ l b/hr----

Ash l b/hr----

Char l b/hr----

Wood l b/hr----

Entha lpy Flow MMBTU/hr (31.8) (29.9) (1.9) (1.8)

118
COMPONENT UNITS 110 112 112H2O 112LIQ 302B 415 415L 415SCRUB 600 618 704W 705W 710 711 713 715 718 FILTCW FILT-H2O ROBRNRCY ROCLEAN RODIRTY SCRB-LIQ SCRB-SLD TOCLARIF

Total Flow lb/hr 937,046 938,445 30,000 28,601 29,562 34,466 35,980 1,514 9,536 101,798 11,729 101,779 24,070 1 25,445 5,864,499 5,864,499 27,940 57,940 23,858 19,084 6,361 57,940 224 58,164

Temperature F 168 153 112 153 130 136 136 136 113 60 110 110 60 60 90 90 110 112 112 66 90 90 112 112 112

Pres sure Psia 15.7 15.7 15.7 15.7 22.7 22.7 22.7 22.7 14.7 14.7 14.7 14.7 14.7 14.7 14.7 74.7 59.7 15.7 15.7 75.0 20.0 14.7 15.7 15.7 15.7

Vapor Fraction - 1.00 1.00 - - - - - - - - - - - - - - - - - - - - - - -

Hydrogen lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Carbon Monoxide lb/hr - - - - 24 - - - - - - - - - - - - - - - - - - 24 24

Nitrogen lb/hr 541,810 541,810 - 0 0 - - - - - - - - - - - - - - - - - - 0 0

Oxygen lb/hr 25,738 25,738 - 0 - - - - - - - - - - - - - - - - - - - 0 0

Argon lb/hr 9,236 9,236 - 0 - - - - - - - - - - - - - - - - - - - 0 0

Carbon Di oxide lb/hr 231,742 231,741 - 0 40 0 0 0 - - - - - - - - - - - - - - - 40 40

Water lb/hr 128,420 129,821 30,000 28,599 29,395 34,464 35,978 1,514 9,536 101,798 11,729 101,779 24,070 1 25,445 5,864,499 5,864,499 27,940 57,940 23,858 19,084 6,361 57,940 54 57,994

Hydrogen Sulfide (H2S) lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

SO2 lb/hr 29 9 - 0 - - - - - - - - - - - - - - - - - - - 0 0

Ammonia (NH3) lb/hr - - - - 15 1 1 0 - - - - - - - - - - - - - - - 15 15

NO2 lb/hr 69 69 - 0 - - - - - - - - - - - - - - - - - - - 0 0

Methane (CH4) lb/hr - - - - 7 - - - - - - - - - - - - - - - - - - 7 7

Ethane (C2H6) lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Ethylene (C2H4) lb/hr - - - - 2 - - - - - - - - - - - - - - - - - - 2 2

Acetylene (C2H2) lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Propane (C3H8) lb/hr - - - - 2 - - - - - - - - - - - - - - - - - - 2 2

I-Butane (C4H10) lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

N-Butane (C4H10) lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Pentane + lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

Benzene (C6H6) lb/hr - - - - 1 - - - - - - - - - - - - - - - - - - 1 1

Tar (C10H8) lb/hr - - - - 18 - - - - - - - - - - - - - - - - - - 18 18

Carbon (Solid) lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

Sulfur (Solid) lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

Olivine + MgO (Solid) lb/hr 2 - - 2 48 - - - - - - - - - - - - - - - - - - 49 49

Methanol lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Ethanol lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Propanol lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

n-Butanol lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

Pentanol+ lb/hr - - - - - - - - - - - - - - - - - - - - - - - - -

Ash lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Char lb/hr 0 - - 0 4 - - - - - - - - - - - - - - - - - - 4 4

Wood lb/hr - - - - 0 - - - - - - - - - - - - - - - - - - 0 0

Enthalpy Flow MMBTU/hr (16.1) (16.2) (2.0) (1.9) (2.0) (2.3) (2.4) (0.1) (0.7) (7.0) (0.8) (6.9) (1.7) (0.0) (1.7) (403.6) (402.3) (1.9) (3.9) (1.6) (1.3) (0.4) (3.9) (0.0) (3.9)

Heat Stream No. MM BTU/hr Work Stream No. HP

QCTOTAL 134.71 A700.WM701 330.3

A700.WP701 635.1

PFD No. Eq. No. Equipment Name

PFD-P820-A304 H-301 Quench Water Reci rcul ation Cooler

PFD-P820-A401 H-411A 1st Stage Water Intercool er

PFD-P820-A401 H-411B 2nd Stage Water Intercooler

PFD-P820-A401 H-411C 3rd Stage Water Intercooler

PFD-P820-A401 H-411D 4th Stage Water Intercooler

PFD-P820-A401 H-411E 5th Stage Water Intercooler

PFD-P820-A402 H-414 Syngas Cool ing Water Exchanger

PFD-P820-A502 H-591 Hi gher Al cohols Product Fi nishi ng Cool er

PFD-P820-A502 H-593 Ethanol Product Fi ni shing Cooler

PFD-P820-A603 H-603 Bl owdown Cool er

PFD-P820-A202 M-201 Sand / Ash Cool er

PFD-P820-A403 to A406 M-401 to M-405 H2S / CO2 Removal & Sulfur Recovery Systems

COOLING WATER USERS

119
Heat Stream No. MM BTU/hr Work Stream No. HP

COMPONENT UNITS 220 221 326 590C 592C

Total Flow l b/hr7 538 12 6,448 51,000

Te mpe ratureF 60 60 60 110 110

Pressure Psi a 31.0 31.0 29.0 35.0 35.0

Va por Fraction 1.00 1.00

Hydrogen l b/hr-----

Ca rbon Monoxide l b/hr-----

Nitrogen l b/hr-----

Oxyge n l b/hr-----

Argon l b/hr-----

Ca rbon Dioxide l b/hr-----

Water l b/hr--- 6140

Hydrogen Sulfi de (H2S) l b/hr0 0 - 0 -

SO2 l b/hr-----

Ammonia (NH3) l b/hr--- 0 -

NO2 l b/hr-----

Methane (CH4) l b/hr-----

Ethane (C2H6) l b/hr-----

Ethyl ene (C2H4) l b/hr--- 0 -

Ace tyl ene (C2H2) l b/hr-----

Propa ne (C3H8) l b/hr--- 0 -

I-Butane (C4H10) l b/hr-----

N-Buta ne (C4H10) l b/hr---00

Pe nta ne + l b/hr-----

Benzene (C6H6) l b/hr---23

Ta r (C10H8) l b/hr--- 0 -

Ca rbon (Sol id) l b/hr-----

Sulfur (Sol id) l b/hr-----

Olivi ne + MgO (Soli d) l b/hr7 538 12 --

Methanol l b/hr---0 256

Ethanol l b/hr--- 515 50,470

Propa nol l b/hr--- 5,258 53

n-Butanol l b/hr--- 7 -

Pe nta nol + l b/hr-----

Ash l b/hr-----

Char l b/hr-----

Wood l b/hr-----

Entha lpy Flow MMBTU/hr (0.0) (0.0) (0.0) (0.2) (1.3)

120
Appendix F. Feedstock Supply Overview from Idaho National 

Laboratory (INL) 

121
 
 

 

 

 

 

 

 

 

 

Title: Woody Biomass Supply Chain Design for Gasification 

2012 Design for NREL Design Report 

Authors: J. Richard Hess, Erin M. Searcy, and Christopher T. Wright 

This section of the process accommodates the delivery of biomass feedstock from the field to the 

biorefinery, short term storage on-site, and the preprocessing/preparation of the feedstock for 

conversion in the gasifier. The design is based on a specific woody biomass feedstock. The feed 

handling and drying sections are shown in PFD-P820-A101 and PFD-P820-A102. 

Standing southern pine trees are cut with an average diameter at breast height (DBH) of 7–10 

inches using a feller buncher with an accumulator arm. Material is harvested 8 hours per day, 5 

days per week, and 50 weeks per year. The cut trees are piled on the ground and transpirationally 

dried from 50% moisture content (MC) to approximately 35% MC, wet basis (for example, see 

Pottie and Guimier 1985, Rummer 2010, Hall 2009, Sinclair 1984). A grapple skidder drags the 

piled trees to the landing, and the trees are piled near the flail chain debarker. The flail chain 

debarker is equipped with a grapple on the end that loads material from the deck into the flail 

chain debarker. About 20% of the tree mass is ejected as bark and tree tops/limbs. The debarked, 

delimbed logs are fed into the chipper, which uses a 2-inch internal screen. Pneumatics may be 

integrated into the chipper, which increases grinder efficiency and provides additional drying 

beyond what would be accomplished without pneumatics. Chips are ejected from the chipper 

directly into a flat floor chip trailer, and once the truck reaches capacity the material is taken to 

the biorefinery. To increase material availability into the system, reduce dry matter losses, and to 

meet U.S. Department of Energy cost targets, some of the limbs are fed through the chipper in a 

stream separate from the clean chips. This slash stream, as we denote it, is passed through a 

trommel screen to sift out high-ash dirt and bark. However, it is assumed that some of the ash 

and bark have already been separated from the limbs and tops during the high-impact flail 

process. The cleaned slash stream is also ejected into the back of a truck and transported to the 

biorefinery. At the biorefinery, the trucks are weighed and a truck dumper unloads the truck into 

a hopper. A dust collection system is present during unloading to prevent excessive dust 

accumulation and limit dry matter losses. The chips are cleaned using an electromagnet, moisture 

is monitored, and material is passed over a vibratory conveyor to remove any excess dirt/ash that 

may still be present. Material is then conveyed into a pile using a circular stacker and overpile 

reclaimer. A five day supply of material is stored at the refinery. The pile, which acts as a queue, 

is on an asphalt pad with a long grate running through it that allows material to flow out of the 

pile, and up to an additional 5% moisture loss results from forced-air ambient air drying. A front 

end loader continuously pushes material onto the grate to maintain flow. A conveyor located at 

the bottom of the pit under the grate conveys the biomass into a waste heat dryer, where the 

biomass is dried to 10% moisture content using hot flue gas from biomass conversion processes. 

From the dryer, the biomass is conveyed into a metering bin where it is fed into the conversion 

process. 

The Idaho National Laboratory (INL), in partnership with other national labs, has developed a 

woody supply system design that is modeled using a woody biomass supply system model. The 

model is being reviewed, and it will be revised to incorporate comments from industry, 

academia, and national labs to improve the accuracy of the model parameters and output. Future 

122
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

model iterations will identify key sensitivities in the system in terms of costs, material flows, and 

equipment performance parameters. The output table from the model showing the costs 

associated with the operations in Figure F-1 is presented in Table F-1. 

Table F-1. Unit Operation Costs for the Design Case Scenario of Gasification of Woody Biomass, 

Using Southern Pine Pulpwood-Size Trees as a Feedstock

Achieving these targets will support DOE biofuels production goals. All values are in U.S. 2007$ per dry 

ton of material. 

Cost ($/dry ton) 

Total feedstock cost through process feed $61.57 

Total feedstock logistics (harvest $46.37 

through insertion to conversion reactor 

inlet) 

Total cost of grower payment $15.20 

Harvest and collection 

Total cost contribution $18.75 

Capital cost contribution $5.60 

Operating cost contribution $13.15 

In-field preprocessing 

Total cost contribution $11.42 

Capital cost contribution $4.20 

Operating cost contribution $7.22 

Transportation and handling 

Total cost contribution $8.95 

Capital cost contribution $2.95 

Operating cost contribution $6.00 

Plant receiving, storage and sueuing, and infeed preprocessing 

Total cost contribution $7.25 

Capital cost contribution $2.10 

Operating cost contribution $5.15 

The total delivered cost of woody biomass from pulpwood-size pine trees extracted from the 

woody biomass supply model is $46.37/dry ton. When the grower payment of $15.20/dry ton is 

added to this, the total feedstock cost through process feed is $61.57/dry ton. This cost is a 

combination of ownership costs and operating costs. 

The model incorporates a combination of values and relationships obtained from other national 

laboratories, publications, consultation with academics and staff from the U.S. Department of 

Agriculture Forest Service, and published and unpublished INL data. Inputs into the model 

include the following. 

123
 
 

Equipment and Building Costs 

• Ownership costs 

o Annual depreciation 

o Interest on the value of the machinery and equipment 

o Property taxes on equipment 

o Insurance 

o Housing (e.g., equipment shed) 

• Operating costs 

o Repair and maintenance 

o Fuel (diesel and electricity) 

o Other materials 

o Labor 

Variables Examined 

• Feedstock variables 

o Biomass yield 

o Biomass removal limit 

• Harvest and collection variables 

o Harvest window 

o Field losses (harvest efficiency) 

o Machine field speed/capacity 

o Machine field efficiency 

o Biomass moisture at harvest (e.g., standing tree moisture) 

o Biomass bulk density (e.g., tree pile or chip density) 

o Distance to landing 

• Storage variables 

o Dry matter loss in storage 

o Machine (e.g., loader) capacity 

• Preprocessing variables 

o Machine capacity 

o Biomass moisture 

o Dry matter losses 

• Handling and transportation variables 

124
 
 

 

 

o Transport distance/winding factor 

o Transporter speed 

o Loader/unloader capacity 

• Plant receiving variables 

o Receiving hours per day 

o Feedstock inventory 

o Feedstock bulk density 

INL will use the model to run sensitivity analyses, examining the impact of changing equipment 

performance parameters and material properties throughout the woody biomass supply chain. 

These analyses will help identify areas on which to focus future research. Ongoing and future 

work on equipment used in the supply chain, order of operations, and material properties will 

identify potential cost reductions achievable with supply-chain improvements and will develop 

feedstock logistics cost targets for future years. 

Woody Biomass Supply System Design 

For the design case scenario, the emphasis is on cost-effective removal of trees of approximately 

7–10 inch diameter. The modeled annual biorefinery capacity is 800,000 dry tons of material and 

all material will be acquired from southern pine pulpwood plantation with a per-acre yield of 30 

green tons at assumed harvest moisture content of 50%. The biomass removal limit is assumed to 

be 95%. 90% of the land is assumed to be forest land, 90% of the production is assumed to be 

pulpwood size trees, and 90% of the producers are assumed to be participating. Figure F-1 

outlines the design for trees on private land. 

125
 
 

 

 

 

 

7 -10" 

Standing Trees 

Ash: 2% 

MC: 50% 

Felling 

Effic: 75% 

Yarding 

Effic: 75% 

Chipper 

Effic: 55% 

Landing 

Delimbing and 

Debarking 

Effic: 70% 

Delimbed, 

Debarked Trees 

Ash: <1% 

MC: 35% 

Debarked Tree 

Chips (2") 

Ash: <1% 

MC: 35% 

Piled Trees 

Ash: 2% 

MC: 50% 

Loader Chipper 

Effic: 40% 

Screened, 

Chipped Slash 

Ash: 1.5% 

MC: 35% 

Piled Slash 

Ash: 8% 

MC: 35% 

Trommel Screen 

Screened Slash 

Ash: 1.5% 

MC: 35% 

2" Chips 

Ash: <1% 

MC: 35% 

Transpirational 

Drying 

Piled, Moisture 

Reduced Trees 

Ash: 2% 

MC: 35% 

Plant Gate 

Circular Stack 

Reclaimer 

Unload/ Handling/ 

Dust Collection 

Chip 

Cleaning 

Even Flow 

Thermochemical 

Conversion – Low 

Pressure 

Low Pressure 

Feedsystem 

Storage/Ambient 

Drying 

Dryer Using Tail 

Gas Combustion 

Biorefinery 

Loader 

Dried Chips 

Ash: <1% 

MC: 10% 

Reactor In-feed 

Conveyor 

Moisture Reduced 

2“ Chips 

Ash: <1% 

MC: 30% 

Truck 

Transport
 

Figure F-1. Feedstock logistics supply system for the design case scenario, using woody biomass as a feedstock. The trees are cut and 

brought to the landing where they are debarked and delimbed, then chipped and sent to the biorefinery for further preprocessing. A 

second material stream consisting of slash (the limbs and tops of trees) is fed through the chipper to decrease total system dry matter 

loss. Note that the ratio of the number of pieces of equipment shown is not 1:1 (i.e., there are, for example, more feller bunchers than 

chippers). The number of pieces of equipment is sized by the operation window and equipment operating capacity. 

126
 
 

 

Appendix F References 

Pottie, M.; Guimier, D. (1985). Preparation of Forest Biomass for Optimal Conversion. Forest 

Engineering Research Institute of Canada International Energy Agency, ISSN 0381-7733. 

Personal communication with Dr. Robert Rummer, Project Leader, Forest Operations Research, 

USDA Forest Service, 521 Devall Drive, Auburn, AL, 36849. 

Hall, P. (2009). Storage Guidelines for Wood Residues for Bioenergy. Prepared by Scion Next 

Energy Biomaterials for EECA (Energy Efficiency and Conservation Authority). 

Sinclair, S.; Hassler, C.; Bolstad, K. (1984). “Moisture Loss in Aspen Logging Residue.” Wood 

and Fiber Science (16:1); pp. 93-96. 

127
 
Appendix G. Synthesis Gas and Char Correlations for Indirect 

Gasifier 

128
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

The gasifier is modeled by correlations based on data from the Battelle Columbus Laboratory 

(BCL) 9 tonne per day test facility. The data and original correlations for the gasifier can be 

found in Bain (1992). Experimental runs were performed for several different wood types 

including red oak, birch, maple, and pine chips; sawdust; and other hard and soft wood chips. 

The original pilot plant data for these runs can be found in Feldmann et al. (1988). The 

temperature range for the data is 1,280°F to 1,857°F and the pressure range is 2.4 to 14.4 psig; 

the majority of the data are in the range of 1,500°F to 1,672°F. 

The BCL test facility’s gas production data was correlated to gasifier temperature with a 

quadratic function in the following form: 

X = a + b T + c T 2 

where the temperature, T, is in units of °F. The coefficients a, b, and c, as well as the units for the 

correlated variable, are shown in Table G-1. Even though there is a correlation for the char 

formation, it is not used; instead the amount and elemental analysis for the char is determined by 

mass differences between the produced syngas and the converted biomass. 

Table G-1. Coefficients for Indirect Gasifier Correlations 

Variable a b c Units 

Dry syngas 28.993 -0.043325 0.000020966 scf gas/lb maf wooda 

CO 133.46 -0.102900 0.000028792 mole % dry gas 

CO2 -9.5251 0.037889 -0.000014927 mole % dry gas 

CH4 -13.82 0.044179 -0.000016167 mole % dry gas 

C2H4 -38.258 0.058435 -0.000019868 mole % dry gas 

C2H6 11.114 -0.011667 0.000003064 mole % dry gas 

H2 17.996 -0.026448 0.000018930 mole % dry gas 

C2H2 -4.3114 0.0054499 -0.000001561 mole % dry gas 

Tar 0.045494 -0.000019759 lb/lb dry wood 

a Scf = standard cubic feet. The standard conditons are 1 atm pressure and 60°F temperature. 

The following general procedure is used for the gasifier production: 

•	 A gasifier temperature T is assumed. 

•	 The mass and molar amounts of carbon, hydrogen, oxygen, sulfur, nitrogen, and ash (as a 

pseudo-element) are determined from the biomass’s ultimate analysis. 

•	 The amount of syngas and its composition is determined from the gasifier correlations. 

•	 The amount of carbon in the syngas and tar is determined. Residual carbon is parsed in 

the char. 

•	 The amount of oxygen in the syngas is determined. A minimum amount of oxygen is 

required to be parsed to the char (4% of biomass oxygen). If there is a deficit of oxygen, 

then the associated water is decomposed to make sure that this amount of oxygen is 

129
 
130 

parsed to the char; if there is excess oxygen, then it is parsed to the char without 

decomposing hydrogen. 

 A set amount of sulfur is parsed to the char (8.3%). All remaining sulfur is set as H2S in 

the syngas. 

 A set amount of nitrogen is parsed to the char (6.6%). All remaining nitrogen is set as 

NH3 in the syngas. 

 The amount of hydrogen in the syngas (including tar, H2S, NH3, and decomposed water) 

is determined. All remaining hydrogen is parsed to the char. 

 All ash is parsed to the char. 

 The heat of formation of the char is estimated from the resulting ultimate analysis from 

this elemental material balance. 

 The gasifier temperature is adjusted so that there is no net heat for an adiabatic reaction. 

The syngas amount and composition will be dependent upon the biomass composition and the 

gasifier temperature. As an example, the resulting syngas composition for the woody biomass 

used in this design report can be seen in Figure G-1. Note from this figure that the amount of 

char decreases with increasing temperature and that the water does not start to decompose until 

high temperatures (here at 1,650°F and higher). 

Figure G-1. Syngas composition for woody biomass used in design report 
Appendix G References 

Bain, R. (January 14, 1992). Material and Energy Balances for Methanol from Biomass Using 

Biomass Gasifiers. 

Feldmann, H.F.; Paisley, M.A.; Applebaus, H.R.; Taylor, D.R. (July 1988). Conversion of Forest 

Residues to A Medium-Rich Gas in a High-Throughput Gasifier. 

131
 
 

 

Appendix H. Equipment Lists and Cost Data for Package 

Systems 

This appendix includes equipment lists and cost data for the following package systems: 

• Gasifier design from Taylor Biomass Energy 

• DEPG acid gas removal system 

• Amine acid gas enrichment system 

• Methanol/H2S absorption system. 

132
 
 

Equipment Included in Gasifier Design from Taylor Biomass Energy 

• Biomass feeding system (excluding storage, pre-treatment, and transport equipment) 

• Gasification process reactors (gasification, combustion, and gas conditioning) 

• All cyclones and interconnecting ducting 

• Air heater 

• Start-up equipment 

• Polishing scrubber (syngas final cleanup and particulate removal) 

• Ash discharge system 

• Steam generation system 

• Emergency flare 

• Air blower 

• Process instrumentation 

• Gasification process air heater 

• Sand makeup system 

• Inert gas system 

• Gasification structural components 

• All to be installed in modules delivered to site (estimated on-site erection 4 to 6 weeks) 

133
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Equipment Costs for DEPG Acid Gas Removal System 

Basis for Cost Estimate 

H2S removal rate 5 lb-mole/h 

CO2 removal rate 1,230 lb-mole/h 

DEPG circulation rate 1,385 gpm 

Absorber operating pressure 3,000 psia 

Regenerator operating pressure 30 psia 

Base Cost Estimate 

Total Purchased Installation Total Source 

Equipment 

Cost (TPEC) 

Factor Installed 

Cost (TIC) 

(2007 k$) (2007 k$) 

DEPG single-stage equipment list 

Absorber pre-cooler 812 3.00 2,433 Doctor et al. 

Absorber column 2,250 2.47 5,558 Vendor quote 

Flash drum 38 1.95 75 Doctor et al. 

Solvent circulation pump 752 1.40 1,056 Doctor et al. 

Lean-rich heat exchanger 519 3.00 1,556 Doctor et al. 

Stripper column 206 2.47 509 Vendor quote 

Stripper reboiler 525 3.00 1,571 Doctor et al. 

Stripper OH condenser 979 3.00 2,933 Doctor et al. 

Stripper OH accumulator 152 1.95 295 Doctor et al. 

Misc. tanks, filters, pumps 400 2.00 800 Aspen Icarus 

DEPG total 6,633* 2.53 16,786* 

* Package system costs correspond to the basis for the estimate. Package costs shown elsewhere in the report are scaled from 

these values based on capacity. 

134
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Equipment Costs for Selective Amine Acid Gas Enrichment System 

Basis for Cost Estimate 

H2S feed rate 5 lb-mole/h 

CO2 feed rate 1,230 lb-mole/h 

CO2 slip Maximize 

Amine circulation rate 200 gpm 

Absorber operating pressure 28 psia 

Regenerator operating pressure 30 psia 

Cost Estimate 

Total Purchased Installation Total Source 

Equipment Cost 

(TPEC) 

Factor Installed 

Cost (TIC) 

(2007 k$) (2007 k$) 

Acid gas enrichment (AGE) equipment list 

Absorber column 145 2.00 290 Vendor quote 

Solution circulation pumps (rich and lean) 39 5.56 217 Aspen Icarus 

Lean-rich heat exchanger 71 3.49 249 Aspen Icarus 

Lean solution trim cooler 115 5.11 588 Aspen Icarus 

Stripper column 170 2.00 340 Vendor quote 

Stripper reboiler 105 1.87 196 Aspen Icarus 

Stripper OH condenser 262 2.51 657 Aspen Icarus 

Stripper OH accumulator 34 5.75 193 Aspen Icarus 

Misc. tanks, filters, pumps 300 2.50 750 Aspen Icarus 

AGE total 1,241* 2.80 3,480* 

* Package system costs correspond to the basis for the estimate. Package costs shown elsewhere in the report are scaled from 

these values based on capacity. 

135
 
 

 

 

 

 

 

 

 

 

 

 

 

 

Equipment Costs for Methanol/H2S Absorption System 

Basis for Cost Estimate 

Maximum H2S recycle rate 20 lb-mole/h 

Maximum methanol recycle 2,000 lb-mole/h 

Absorber operating pressure 30 psia 

Cost Estimate 

Methanol absorption equipment list 

Total Purchased 

Equipment Cost 

(TPEC) 

(2007 k$) 

Installation 

Factor 

Total 

Installed 

Cost (TIC) 

(2007 k$) 

Source 

Absorber column 45 3.36 150 Icarus 

Methanol pump 76 1.86 141 Icarus 

Lean-rich heat exchanger 33 3.31 110 Icarus 

Methanol air cooler 214 1.40 300 Icarus 

Methanol refrigeration chiller 61 2.55 155 Icarus 

Total 428* 2.00 856* 

* Package system costs correspond to the basis for the estimate. Package costs shown elsewhere in the report are scaled from 

these values based on capacity. 

136
 
Appendix I. Technology and Cost Targets for Gasification, Tar 

Reforming, and Alcohol Synthesis 

137
 
 

 

 

 

 

 

 

 

 

 

 

Unit Operation Cost Contribution Estimates (2007$) and Technical Projections for Thermochemical Conversion to Ethanol Baseline 

Process Concept
 

(Process Concept: Woody Energy Crop, Gasification, Gas Cleanup, Mixed Alcohol Synthesis, Ethanol Recovery and Purification) 

Processing Area Cost Contributions and 

Key Technical Parameters Metric 2007 SOT a 2008 SOT 2009 SOT 2010 SOT 2011 

Projection 

2012 

Projection 

Process Concept: Gasification, Syngas 

Cleanup, Mixed Alcohol Synthesis & Recovery 

Woody 

Feedstock 

Woody 

Feedstock 

Woody 

Feedstock 

Woody 

Feedstock 

Woody 

Feedstock 

Woody 

Feedstock 

Conversion Contribution $/gal EtOH $3.35 $2.11 $2.03 $1.65 $1.62 $1.31 

Year $ basis 2007 2007 2007 2007 2007 2007 

EIA reference case b $/GGE c $2.18 $2.57 $1.69 $2.29 $2.47 $2.62 

$/gal EtOH $1.46 $1.72 $1.13 $1.53 $1.66 $1.76 

Projected minimum ethanol selling price d $/gal EtOH $4.75 $3.35 $3.26 $2.70 $2.51 $2.05 

Total project investment per annual gallon $ $12.76 $9.47 $9.24 $7.96 $7.85 $7.60 

Plant capacity (dry feedstock basis) tonnes/day 2,000 2,000 2,000 2,000 2,000 2,000 

Ethanol yield gal EtOH/dry ton 62 70 70 79 80 84 

Mixed alcohol yield gal MA/dry ton 67 77 78 88 89 94 

Feedstock 

Total cost contribution $/gal EtOH $1.40 $1.24 $1.22 $1.05 $0.90 $0.73 

Capital cost contribution $/gal EtOH - - - - - -

Operating cost contribution $/gal EtOH $1.40 $1.24 $1.22 $1.05 $0.90 $0.73 

Feedstock cost $/dry U.S. ton $86.25 $86.25 $86.25 $82.70 $71.60 $61.57 

Feedstock moisture at plant gate wt % H2O 50% 50% 50% 40% 40% 30% 

In-plant handling and drying $/dry U.S. ton $22.65 $22.65 $22.65 $20.60 $14.30 $7.25 

Cost contribution $/gal EtOH $0.37 $0.32 $0.32 $0.26 $0.18 $0.09 

Feed moisture content to gasifier wt % H2O 10% 10% 10% 10% 10% 10% 

Energy content (LHV, dry basis) Btu/lb 8,000 8,000 8,000 8,000 8,000 8,000 

Gasification 

Total cost contribution $/gal EtOH $0.37 $0.33 $0.33 $0.29 $0.29 $0.28 

Capital cost contribution $/gal EtOH $0.21 $0.19 $0.19 $0.17 $0.16 $0.16 

Operating cost contribution $/gal EtOH $0.16 $0.14 $0.14 $0.13 $0.13 $0.12 

Raw dry syngas yield lb/lb dry feed 0.78 0.78 0.78 0.78 0.78 0.78 

Raw syngas methane (dry basis) mole % 15% 15% 15% 15% 15% 15% 

Gasifier efficiency (LHV) % LHV 74% 74% 74% 74% 74% 74% 

138
 
 

 

 

 

 

 

 

Processing Area Cost Contributions and 

Key Technical Parameters Metric 2007 SOT 2008 SOT 2009 SOT 2010 SOT 2011 

Projection 

2012 

Projection 

Synthesis gas cleanup (reforming and quench) 

Total cost contribution $/gal EtOH $1.22 $0.61 $0.58 $0.42 $0.43 $0.17 

Capital cost contribution $/gal EtOH $0.14 $0.12 $0.12 $0.10 $0.10 $0.10 

Operating cost contribution $/gal EtOH $1.07 $0.49 $0.46 $0.32 $0.33 $0.07 

Tar reformer (TR) exit CH4 (dry basis) mole % 13% 5% 4% 2% 2% 2% 

TR CH4 conversion % 20% 50% 56% 80% 80% 80% 

TR benzene conversion % 80% 98% 98% 99% 99% 99% 

TR tars conversion % 97% 97% 97% 99% 99% 99% 

Catalyst replacement % inventory/day 1.0% 1.0% 1.0% 1.0% 1.0% 0.1% 

Acid gas and sulfur removal 

Total cost contribution $/gal EtOH $0.27 $0.21 $0.20 $0.17 $0.17 $0.17 

Capital cost contribution $/gal EtOH $0.17 $0.13 $0.12 $0.11 $0.11 $0.10 

Operating cost contribution $/gal EtOH $0.10 $0.08 $0.08 $0.07 $0.06 $0.06 

Sulfur level at reactor inlet (as H2S) ppmv 70 70 70 70 70 70 

Synthesis gas compression and power recovery expansion 

Total cost contribution $/gal EtOH $1.28 $0.84 $0.81 $0.67 $0.67 $0.67 

Capital cost contribution $/gal EtOH $0.65 $0.39 $0.37 $0.29 $0.30 $0.29 

Operating cost contribution $/gal EtOH $0.63 $0.45 $0.44 $0.38 $0.38 $0.38 

Electricity production from syngas expander 

(credit included in operating cost above) $/gal EtOH ($0.35) ($0.15) ($0.14) ($0.08) ($0.09) ($0.09) 

Fuels synthesis reaction 

Total cost contribution $/gal EtOH $0.24 $0.12 $0.11 $0.06 $0.04 $0.03 

Capital cost contribution $/gal EtOH $0.24 $0.19 $0.18 $0.16 $0.16 $0.15 

Operating cost contribution $/gal EtOH $0.00 ($0.07) ($0.08) ($0.10) ($0.12) ($0.12) 

Pressure psia 3,000 3,000 3,000 3,000 3,000 3,000 

Single pass CO conversion % CO 25% 24% 25% 26% 29% 29% 

Overall CO conversion % CO 55% 68% 70% 80% 79% 79% 

Selectivity to alcohols % CO (CO2 free) 78% 81% 81% 81% 81% 81% 

Selectivity to ethanol % CO (CO2 free) 59% 63% 63% 63% 63% 63% 

Ethanol productivity g/kg-cat/h 101 128 132 143 153 160 

Mixed alcohols co-product 

(credit included in operating cost above) $/gal EtOH ($0.18) ($0.22) ($0.22) ($0.23) ($0.24) ($0.24) 

139
 
 

 

 

Processing Area Cost Contributions and 

Key Technical Parameters Metric 2007 SOT 2008 SOT 2009 SOT 2010 SOT 2011 

Projection 

2012 

Projection 

Product recovery and purification 

Total cost contribution $/gal EtOH $0.14 $0.12 $0.12 $0.11 $0.11 $0.10 

Capital cost contribution $/gal EtOH $0.10 $0.09 $0.08 $0.08 $0.08 $0.07 

Operating cost contribution $/gal EtOH $0.04 $0.04 $0.04 $0.03 $0.03 $0.03 

Balance of plant 

Total cost contribution $/gal EtOH ($0.17) ($0.12) ($0.11) ($0.09) ($0.09) ($0.10) 

Capital cost contribution $/gal EtOH $0.30 $0.24 $0.23 $0.21 $0.21 $0.20 

Operating cost contribution $/gal EtOH ($0.47) ($0.35) ($0.35) ($0.31) ($0.30) ($0.30) 

Electricity production from steam turbine 

(credit included in operating cost above) $/gal EtOH ($0.60) ($0.46) ($0.45) ($0.40) ($0.40) ($0.39) 

Electricity production kWh/gal EtOH 16.6 10.7 10.3 8.5 8.5 8.4 

Electricity consumption (entire process) kWh/gal EtOH 16.6 10.7 10.3 8.5 8.5 8.4 

Water consumption e gal H2O/gal EtOH 7.0 3.7 3.5 2.8 2.7 2.6 

Fuel ethanol case reference (model run # ) 

AD-FY07-R236­

50pctMoistureV24.xls 

AD-FY08-R236­

50pctMoistureV24.xls 

AD-FY09-R236­

50pctMoistureV24.xls 

AD-FY10-R236­

40pctMoistureV24.xls 

AD-FY11-R236­

40pctMoistureV24.xls R236-V24.xls 

a SOT: State of Technology 

b EIA, “Annual Energy Outlook 2009,” Table 112, U.S., http://www.eia.doe.gov/oiaf/archive/aeo09/supplement/suptab_112.xls 

c Using a 0.67 gallon gasoline per gallon ethanol conversion factor 

d Conceptual design result with expected margin of error ±30% 

e 2012 water consumption of 2.6 gallons of water per gallon of ethanol represents the base case value. The water consumption value resulting from an optimized water treatment and 

recycle design is 2.0 gallons of water per gallon of ethanol. 

140
 
Appendix J. Fuel-Grade Ethanol Specifications
 

141
 
 

 

 

 

 

 

Fuel‐Grade Ethanol Specifications 

 Below are the primary quality specifications for denatured fuel ethanol for blending with gasoline 

meeting Federal requirements. The state of California has additional restrictions that apply in 

addition to the performance requirements in ASTM D 4806. 

 Specifications Contained in ASTM D 4806 Standard Specification for
 

Denatured Fuel Ethanol for Blending with Gasoline

Property Specification ASTM Test Method 

Ethanol Content (Volume %) 92.1 Minimum D 5501 

Methanol Content (Volume %) 0.50 Maximum D 5501 

Solvent-Washed Gum (mg/100 ml) 5.0 Maximum D 381 

Water Content (Volume %) 1.0 Maximum E 203 

Denaturant Content (Volume %) 1.96 Minimum D 4806 

4.76 Maximum 

Inorganic Chloride Content (mg/L) 40 Maximum D 512 

Copper Content (mg/kg) 0.10 Maximum D1688 

Acidity (Mass % as acetic acid CH3COOH) 0.0070 Maximum D1613 

pHe 6.5 Minimum D 6423 

9.0 Maximum 

Appearance Visibly free of suspended or precipitated 

contaminants (clear & bright) 

Source: 

Renewable Fuels Association, Industry Guidelines, Specifications, and Procedures. 

http://www.ethanolrfa.org/pages/industry-resources-guidelines 

Note: 

ASTM = American Society for Testing and Materials

142
Appendix K. Direct Costs Included in Equipment Installation 

Factors from Harris Group Inc. 

143
 
 

 

 

 

 

 

 

 

 

Direct Costs Included in Installation Factors from Harris Group Inc. 

Site Preparation/Civil 

1.	 General site preparation (grading, cut, and backfill) 

2.	 Excavation and backfill requirements related to equipment foundations. Generally, the 

excavated and backfill volume was estimated to be roughly 80%–90% of the in-ground 

concrete volume. 

3.	 Underground piping 

4.	 Excavation and backfill requirements for sanitary, storm water, firewater and process 

sewer culverts, piping, etc. 

5.	 Excavation and backfill requirements for retention/drainage ponds, if necessary 

6.	 Any other required cut/backfill 

7.	 Any required paving and gravel including proper drainage (curbs and gutters, etc.) 

A. Access 

B. Parking 

C. Truck staging/truck scales 

D. Roads 

8.	 Any required sumps and pumps for potentially contaminated waste streams or streams to 

be recycled back to the process 

9.	 Containment berms 

10. Any required retention ponds or in-ground water storage 

Demolition 

1.	 If not a green field project an allowance must be included for removal of the existing 

concrete, steel, equipment, etc. 

Concrete/Foundations 

1.	 Depending on soil quality the foundation design and requirement for piles can be
 

determined. Typically we assume an unknown soil quality to be average, having 

approximately a 3000–4000 psf soil bearing pressure.
 

A. Any requied piles and pile caps. Pile length determined from soil report. 

B. Equipment foundations that are optimized for the arrangement of building 

columns and actual loads 

C. Containment 

D. All foundations are priced including rebar, rather than mesh, form work, hardware 

(anchor bolts, iron, etc.), concrete, finishing, and form stripping 

2.	 Bulk storage slabs (chip piles, etc.) 

144
 
 

 

Buildings 

1.	 Possible necessary buildings include the following 

A. Electrical building (switchgear, MCC, and rack room) 

B. Control room 

C. Maintenance shops 

D. Spares and parts storage 

E. Continuous emission monitoring (CEMS) building to house the monitoring and 

control equipment 

F.	 Lockers/restrooms 

G. Office space 

H. Meeting space 

I.	 Guard shack 

2.	 Main process area can include the following 

A. Cranes 

B. Access doors, hatches, etc. 

3.	 Elevators for multi-floor structures 

Structural 

1.	 Equipment support steel and hardware (assume an average wind and seismic zone for 

unknown locations) 

2.	 Pipe bridges and miscellaneous support steel 

3.	 Equipment access steel 

A. Stairs 

B. Grating/decking 

C. Ladders 

D. Railings 

E. Gates 

4.	 Structural steel for cranes 

5.	 Maintenance monorails 

Mechanical 

1.	 At times the following equipment and/or systems can be included in the installation 

factors for preliminary/feasibility estimates 

A. Process air systems 

145
 
 

 

 

 

B. Instrument air systems 

C. Wastewater treatment systems 

D. Cooling water systems 

E. Steam generating/water treatment systems 

Pumps, Piping, Ductwork, and Valves 

1. All pumps not included with equipment 

2. All process and utility piping 

A. Including all valves, flanges, fittings, and hardware not included with equipment 

3. All process and utility ductwork 

A. Including all dampers, flanges, fittings, and hardware 

4. Miscellaneous storm drainage, sanitary sewer and process sewer piping 

5. Utility piping runs to include connections, back flow prevention, etc. at sources 

A. Process water (if not produced on-site) 

B. Cooling water (if not produced on-site) 

C. Process air (if not produced on-site) 

Electrical 

1. High voltage breaker and feeder located in a high voltage bus room 

2. Medium voltage feeder to loop feed the unit substations 

3. Unit substations 

A. Outdoor oil-filled transformer(s) with close-coupled primary switch 

B. Indoor secondary switchgear 

C. Cable and conduit for the power distribution feeders between the outdoor 

transformers and the indoor switchgear and between the switchgear and the 

MCCs is included 

4. MCCs 

A. Low voltage motor control centers will contain variable frequency drives (VFDs) 

up to 250 HP. Low voltage VF drives larger than 250 HP will be free standing 

drives separately fed from a feeder breaker in the unit substation’s 600 volt 

secondary switchgear. All VF drives will be furnished with input line reactor (to 

limit harmonics to upstream equipment) and output filters (to clamp voltage 

spikes and limit harmonics to the motors). 

B. All VFDs larger than 400 HP will be medium voltage and will be separately fed 

from a feeder breaker in the unit substation’s 5 kV secondary switchgear 

C. Motors not included with equipment 

146
 
 

 

 

 

 

 

5.	 Motor wiring 

A. Motor lead wiring to include cable, conduit, and terminations 

B. Motor lead wiring for VFDs shall be continuous corrugated aluminum sheath 

CLX type cable with three symmetrical grounding conductors installed in ladder 

type cable tray 

6.	 Instrument wiring (including conduit) 

7.	 Uninterrupted power supply (UPS) for control systems and any critical systems such as 

burner management systems, etc. 

8.	 Grounding 

9.	 Lightning protection 

10. Lighting 

11. Electrical building (switchgear, MCC, and rack room) 

12. Cable trays 

Instrumentation and Controls 

1.	 Either DCS or PLC system (chassis, power supply, processors, etc.) 

2.	 Operator stations (HMI software and PCs) 

3.	 Hardware (I/O modules, etc.) 

4.	 Field instruments and transducers are 4–20 mA type with twisted shield pair wiring, 

discrete devices are normally 120 VAC. Field bus? 

5.	 Control/actuated valves 

6.	 PLC programming and programming software 

7.	 DCS interface to PLC(s) (if DCS is utilized) 

8.	 DCS programming (if DCS is utilized) 

9.	 Any statistics, data archiving, or other control related software. 

Fire Protection/Safety 

1.	 Fire Protection 

A. Above- and below-ground piping, valves, fittings, flanges, hardware, etc. 

B. Hydrants and sprinklers 

2.	 Safety 

A. Access stairs, ladder, etc. not included with structural steel 

B. Eye wash and shower stations, including piping, valves, etc. 

147
 
Painting 

1. Allowance for painting structural steel and equipment 

Insulation/Refractory/Heat Tracing 

1. Piping insulation for heat conservation and/or personnel safety 

2. Any refractory not included with equipment 

3. Any heat tracing required for cold weather service 

House Keeping 

1. Wash up hose stations 

148
 
 

Appendix L. Plant Footprint and Land Cost Calculation from 

Harris Group Inc. 

149
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Plant Footprint and Land Cost Calculation from Harris Group Inc.

 Basis: 

2,000 dry metric tons/day feedstock basis or 2,205 dry 

short tons/day or 83.7 gallons per dry short ton of feedstock 

or ~64.6 MM gal/year at ~ 350 day/year of operation 

Estimated Footprint Calculated Area 

Width Length

 Areas ft ft ft^2 Acre 

1 Truck staging, scales, scale shack, guard shack etc. 250 500 125,000 2.87 

2 Feedstock handling and storage - x2-truck dumps/hoppers, x2 800 1,000 800,000 18.37 

scalping screens/hogs, x2-400’ diameter stacker reclaimers and 

a bunch of conveyors 

3 Feedstock drying and storage - x2-dryers/WESPs/RTOs and x2- 200 500 100,000 2.30 

dry feedstock silos 

4 Gasifier Island - x2-gasifiers, x2-tar reformers, bed/sorbent 125 150 18,750 0.43 

media handling w/ silos, ash handling w/ silo(s), instrument air 

system etc. 

5 Syngas scrubbing, compression, synthesis, heat exchangers, 150 150 22,500 0.52 

expander etc. 

6 Acid gas removal, acid gas enrichment, sulfur recovery 150 150 22,500 0.52 

7 Alcohol purification (distil, mole sieve etc.) 100 150 15,000 0.34 

8 Waste water treatment, boiler water treatment, condensate 200 200 40,000 0.92 

treatment, DA, steam distribution, power generation etc. 

9 Cooling towers, pumps etc. 100 150 15,000 0.34 

10 Tank Farm - product tanks, fuel, chemicals, fire water, process 500 500 250,000 5.74 

water etc. 

11 Air separation plant, compressors, tanks etc. 50 50 2,500 0.06 

12 Storm water system - retention pond etc. 200 200 40,000 0.92 

13 Electrical feed, transformers etc. 100 100 10,000 0.23 

14 Electrical Building 25 100 2,500 0.06 

15 Control Room 25 50 1,250 0.03 

16 Offices, etc. 100 100 10,000 0.23 

17 Maintenance shop, maintenance warehouse/yard 200 200 40,000 0.92 

18 Parking, roads, etc. 500 500 250,000 5.74 

19 Buffer 871,200 20.00 

20 Unusable land (topography and environmentally) 1,306,800 30.00 

21 Other (10% of process areas and buildings) 151,500 3.48 

22 Easement (10% of all above) 409,450 9.40 

23 Contingency (10% of all above) 450,395 10.34

 Total 4,954,345 113.74

 Specific Land Cost $ / Acre 14,000

 Site area Acre 113.74

 Total Land Cost $ 1,592,306 

150
Appendix M. Pinch Analysis and Heat Integration System Design
 

151
 
Equip ID Duty ID Service Duty 

(MMBTU/Hr) 

Hot In T 

(Deg F) 

Hot Out T 

(Deg F) 

Cold In T 

(Deg F) 

Cold Out T 

(Deg F) 

LMTD 

(Btu/Hr

H-201 QH201 Syngas Quench Inlet / Combustion Air Pre-Heat Exchanger * 4.6 331.6 290.8 277.1 316.3 14.5 2H-202 QH202 Tar Reformer Effluent / Combustion Air Pre-Heat Exchanger * 9.9 436.7 331.6 316.3 400.0 24.4 2

H-302 QH302 Steam Blowdown / Unreacted Syngas Exchanger 4.4 581.0 160.0 72.3 564.1 43.0 3

H-303 QH303 Syngas Quench Inlet / Make-Up Water Exchanger 7.2 254.4 160.2 145.0 239.4 15.1 3

H-304 QH304 Syngas Quench Inlet / Mol-Sieve Pre-Heat Exchanger 5.7 331.6 290.8 200.4 237.9 92.1 2

H-305 QH305 Alcohol Synthesis Effluent / Combustion Air Pre-Heat Exchanger 1.6 331.6 257.3 160.9 316.3 44.0 4H-306 QH306 Syngas Quench Inlet / Combustion Air Pre-Heat Exchanger 2.6 331.6 257.3 160.9 316.3 44.0 4

H-307 QH307 Syngas Quench Inlet / Combustion Air Pre-Heat Exchanger 5.9 331.6 254.4 160.9 316.3 43.2 4

H-308 QH308 Tar Reformer Effluent / Combustion Air Pre-Heat Exchanger 5.3 436.7 331.6 316.3 400.0 24.4 2

H-309 QH309 Tar Reformer Effluent / Combustion Air Pre-Heat Exchanger 12.7 865.8 787.3 400.0 600.0 322.7 2

H-310 QH310 Catalyst Regenerator Flue Gas / Tar Reformer Steam & Recycle Exchanger 69.9 1,776.8 1,139.3 575.1 1,700.0 244.4 2

H-311 QH311 Tar Reformer Effluent / Unreacted Syngas Exchanger 9.1 1,670.0 865.8 577.5 1,139.0 397.4 2

H-401 QH401 Alcohol Synthesis Feed / Effluent Exchanger 127.7 611.3 331.6 316.3 556.9 30.8 9

H-402 QH402 Alcohol Synthesis Feed / Effluent Exchanger 73.6 331.6 200.0 182.1 316.3 16.6 9

H-403 QH403 Compressor Interstage / Deaerator Feed Exchanger 25.0 388.1 249.9 235.0 346.7 25.9 3

H-404 QH404 Compressor Interstage / Deaerator Feed Exchanger 15.3 359.6 249.9 235.0 346.7 13.9 5

H-405 QH405 Compressor Interstage / Deaerator Feed Exchanger 15.5 358.7 249.9 235.0 346.7 13.4 10

H-406 QH406 Compressor Interstage / Mol-Sieve Pre-Heat Exchanger 16.0 249.9 170.0 66.8 237.9 42.3 3

H-407 QH407 Compressor Interstage / Mol-Sieve Pre-Heat Exchanger 11.1 249.9 170.0 66.8 237.9 42.3 5

H-408 QH408 Compressor Interstage / Mol-Sieve Pre-Heat Exchanger 11.4 249.9 170.0 66.8 237.9 42.3 7

H-409 QH409 Alcohol Synthesis Effluent / LO-CAT Pre-Heat Exchanger 0.0 200.0 176.1 80.0 150.0 70.6 5H-415 QH415 Alcohol Synthesis Effluent / Syngas Recycle Exchanger 19.6 200.0 175.0 115.2 180.0 36.3 9H-416 QH416 Alcohol Synthesis Effluent / CO2-Rich Gas Exchanger 1.0 200.0 176.1 94.0 160.2 58.5 8

H-417 QH417 Alcohol Synthesis Effluent / Mol-Sieve Pre-Heat Exchanger 5.4 220.0 200.0 66.8 200.4 59.3 2

H-418 QH418 Alcohol Synthesis Effluent / Expander Inlet Exchanger 9.7 331.6 257.3 95.4 316.3 62.1 8

H-419 QH419 Tar Reformer Effluent / Expander Inlet Exchanger 3.8 436.7 331.6 316.3 400.0 24.4 3

H-420 QH420 Tar Reformer Effluent / Expander Inlet Exchanger 25.1 1,670.0 865.8 400.0 950.0 583.7 3

H-421 QH421 Alcohol Synthesis Effluent / Expander Inlet Exchanger 4.3 331.6 257.3 30.0 316.3 78.6 7

H-422 QH422 Tar Reformer Effluent / Expander Inlet Exchanger 11.0 1,670.0 865.8 316.3 950.0 630.9 2H-423 QH423 Tar Reformer Effluent / Expander Inlet Exchanger 16.3 1,670.0 865.8 594.5 950.0 459.7 2

H-424 QH424 Alcohol Synthesis Effluent / CO2-Rich Gas Exchanger 2.4 331.6 257.3 160.2 316.3 44.3 4

H-425 QH425 Char Combustor Flue Gas / CO2-Rich Gas Exchanger 4.5 1,808.2 787.3 316.3 1,170.0 550.4 2

152
Equip ID Duty ID Service Duty 

(MMBTU/Hr) 

Hot In T 

(Deg F) 

Hot Out T 

(Deg F) 

Cold In T 

(Deg F) 

Cold Out T 

(Deg F) 

LMTD 

(Btu/Hr

H-426 QH426 Tar Reformer Effluent / CO2-Rich Gas Exchanger 10.1 1,670.0 865.8 316.3 1,170.0 524.3 2

H-427 QH427 Catalyst Regenerator Flue Gas / Alcohol Synthesis Feed Exchanger 20.4 1,139.3 787.3 556.9 595.4 365.0 4

H-501 QH501 Ethanol Product / Mol-Sieve Pre-Heat Exchanger 4.8 229.5 120.0 66.8 214.5 30.2 9

H-502 QH502 Higher Alcohols Product / Mol-Sieve Pre-Heat Exchanger 0.8 264.8 120.0 66.8 214.5 51.8 9

H-503 QH503 Syngas Quench Inlet / Water-Methanol Recycle Exchanger 12.1 331.6 160.2 140.0 316.3 17.6 3H-506 QH506 Syngas Quench Inlet / Mol-Sieve Pre-Heat Exchanger 10.4 220.0 160.2 66.8 200.4 47.3 2H-507 QH507 Alcohol Synthesis Effluent / Mol-Sieve Pre-Heat Exchanger 1.0 331.6 257.3 237.9 257.9 40.7 5

H-508 QH508 Alcohol Synthesis Effluent / Mol-Sieve Pre-Heat Exchanger 0.9 331.6 257.3 214.5 237.9 65.0 4

H-509 QH509 Catalyst Regenerator Flue Gas / Water-Methanol Recycle Exchanger 0.5 1,139.3 787.3 316.3 590.0 509.1 2

H-510 QH510 Tar Reformer Effluent / Water-Methanol Recycle Exchanger 1.5 865.8 787.3 316.3 590.0 364.7 2

H-602 QH602 Steam Turbine Exhaust / Syngas Recycle Exchanger 3.9 234.1 160.0 115.2 180.0 49.3 4

H-604 QH604 Steam Turbine Exhaust / Mol-Sieve Pre-Heat Exchanger 4.4 234.1 195.0 66.8 200.4 70.8 3

H-605 QH605 Syngas Quench Inlet / Make-Up Water Exchanger 6.5 331.6 160.2 60.0 145.0 139.0 3

H-606 QH606 Syngas Quench Inlet / BFW Pre-Heat Exchanger 6.5 257.3 220.0 198.0 232.3 23.5 3

H-607 QH607 Alcohol Synthesis Effluent / BFW Pre-Heat Exchanger 10.0 257.3 220.0 198.0 232.3 23.5 3H-608 QH608 Tar Reformer Effluent / BFW Pre-Heat Exchanger 63.7 787.3 436.7 346.7 582.0 139.8 3

H-609 QH609 Catalyst Regenerator Flue Gas / BFW Pre-Heat Exchanger 26.1 787.3 549.4 346.7 582.0 204.0 3

H-610 QH610 Char Combustor Flue Gas / BFW Pre-Heat Exchanger 38.1 787.3 532.0 346.7 582.0 195.1 3

H-611 QH611 Catalyst Regenerator Flue Gas Steam Generator 17.7 1,139.3 787.3 581.0 581.0 353.6 3

H-612 QH612 Char Combustor Flue Gas Steam Superheater 147.9 1,808.2 787.3 580.0 1,000.0 441.6 3

H-613 QH613 Tar Reformer Effluent Steam Superheater 74.6 1,670.0 865.8 698.1 1,000.0 362.6 2

Total 999.6 

* Costs for H-201 and H-202 are included in total capital for the gasifier and tar reformer based on quotation provided by Taylor Biomass Energy.

** Heat transfer coefficients used for sizing equipment in the heat integration network are from the GPSA Engineering Data Book, FPS Version, Volume I, Section 9. Eleventh Edit

153
Appendix N. Summary of Design Report Assumptions
 

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The table below presents a summary of assumptions for the thermochemical ethanol design compared to the assumptions for the 

previous design version (2007) and the updated biochemical ethanol design. Sources of justification for changes in assumptions are 

experience and feedback from Harris Group Inc., feedback from peer reviewers, current financial atmosphere and markets, feedback 

from DOE, and inclusion of DOE programs such as loan guarantees. 

Parameter 2011 Thermochemical 

Ethanol Design Report 

2011 Biochemical 

Ethanol Design Report 

2007 Thermochemical 

Ethanol Design Report 

Cost year for analysis 2007 2007 2005 

Plant financing by equity / debt 40% / 60% 40% / 60% 100% / 0% 

Internal rate of return (IRR) 10% after-tax 10% after-tax 10% after-tax 

Term for debt financing 10 years 10 years N/A 

Interest rate for debt financing 8% 8% N/A 

Plant life / analysis period 30 years 30 years 20 years 

Depreciation method 

MACRS 

7 years general plant 

7 years steam/power 

(no export of electricity) 

MACRS 

7 years general plant 

15 years steam/power 

(export of electricity) 

MACRS 

7 years general plant 

20 years steam/power 

Income tax rate 35% 35% 39% 

Plant construction cost schedule 3 years 

(8% Y1, 60% Y2, 32% Y3) 

3 years 

(8% Y1, 60% Y2, 32% Y3) 

2.5 years 

(8% Y1, 60% Y2, 32% Y3) 

Plant decommissioning costs $0 $0 $0 

Plant salvage value $0 $0 $0 

Start-up period 3 months 3 months 6 months 

Revenue and costs during start-up 

Revenue = 50% of normal 

Variable costs = 75% of normal 

Fixed costs = 100% of normal 

Revenue = 50% of normal 

Variable costs = 75% of normal 

Fixed costs = 100% of normal 

Revenue = 50% of normal 

Variable costs = 75% of normal 

Fixed costs = 100% of normal 

155
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Parameter 2011 Thermochemical 

Ethanol Design Report 

2011 Biochemical 

Ethanol Design Report 

2007 Thermochemical 

Ethanol Design Report 

On-stream percentage 96% 

(8,410 operating hours / year) 

96% 

(8,410 operating hours / year) 

~96% 

(8,406 operating hours / year) 

Land purchase cost $1.61MM 

(115 acres at $14,000 per acre) 

$1.85MM 

(132 acres at $14,000 per acre) 

$3.34MM 

(6% of TPEC) 

Equipment installation factor 

(TIC / TPEC) 1.98 1.50 2.47 

Warehouse costs (direct) $0 4.0% of ISBL total installed costs $0 

Site development costs (direct) 4.0% of ISBL total installed costs 9.0% of ISBL total installed costs $0 

Additional piping (direct) $0 4.5% of ISBL total installed costs $0 

Indirect costs 60% of total direct costs 60% of total direct costs 39% of total direct costs 

Working capital cost 5% of fixed capital investment 5% of fixed capital investment 5% of fixed capital investment 

FCI / TPEC 3.45 2.59 3.43 

Annual plant maintenance costs 3% of total fixed capital investment 3% of ISBL fixed capital 

investment 2% of fixed capital investment 

Annual insurance and taxes 0.7% of fixed capital investment 0.7% of fixed capital investment 2% of fixed capital investment 

Feedstock basis Woody biomass (southern pine) Corn stover Woody biomass (hybrid poplar) 

Feedstock cost 

$61.57 per dry ton including 

capital and operating costs up to 

inlet of gasifier reactor 

$58.50 per dry ton including 

capital and operating costs up to 

inlet of pretreatment 

$35.00 per dry ton 

Alcohol synthesis yield basis Kinetic model from 

The Dow Chemical Company N/A Stoichiometric model 

156
 
Appendix O. Reviewer Comments on Draft Design Report and 

Responses 

157
 
 

 

 

 

 

 

 

 

 

 

 

 

 

This appendix presents a summary of reviewer comments from the draft version of the report, 

which was issued for peer review in December 2010. Each comment from the peer review panel 

is summarized below and followed by a response from NREL authors (and INL authors for Area 

100 comments). The responses describe any actions taken to address the comments in the final 

draft of the report. 

Area A100: Feed Handling and Drying 

1.	 Comment: It is understood that power cost for feed handling is included in the INL 

feedstock cost. However, as a practical matter, that power needs to come from elsewhere in 

the plant. This means more syngas converted to power and less ethanol production than the 

model would otherwise predict. It’s unclear from the report whether this allowance has been 

made. 

Response: This allowance has not been made in the model with exception of the heat from 

flue gas required to dry the feedstock to target moisture content (from 30 wt % to 10 wt % in 

the base case design). However it can be said that the economic impact from increased 

syngas consumption for feedstock processing (electricity and other energy) will be offset by 

reduced feedstock cost. This is because the feedstock cost in the model already includes 

allowances for all capital and operating costs (including energy) associated with feedstock 

supply to the throat of the gasifier. 

2.	 Comment: Though not modeled explicitly, the dryer exhaust is likely to contain VOCs 

(volatile organic compounds). I doubt you can just vent it to atmosphere. Perhaps this could 

be routed to combustors downstream? 

Response: A flue gas scrubbing system was added to the design to control SO2 and 

particulate emissions from the biorefinery flue gas. NREL and INL will assess additional 

environmental requirements in future design updates when data is available for VOC 

components in the dryer outlet flue gas. 

3.	 Comment: Starting at 35% moisture is not realistic as we've done lots of feedstock work and 

you can't expect it to dry from a 50% moisture tree (just cut) to 35% moisture in any short 

period of time. It will cost too much to leave it there for a long time (sometimes months) and 

reclaim it later. Expecting it to dry to 30% is also not realistic by chipping and transporting in 

a truck. Unless it is spread out in the sun it doesn't dry by itself. So in the end, one has to dry 

it from 50% to 10%, which would cost more in capital and operating. Possibly close to 

$10/dry ton just for these operations. 

Response: Idaho National Laboratory (INL) has deemed the feedstock technical and cost 

targets presented in the report to be achievable by 2012. INL will release a report on woody 

feedstocks that will include more detail on processing methods and assumptions. The 

sensitivity study in Figure 16 (case 8) shows the impact on the MESP if feed is supplied with 

50 wt % moisture at a cost of $61.57/dry ton. 

4.	 Comment: We don’t necessarily disagree with the per ton biomass costs but it is noticeably 

different than before. A footnote to highlight the impacts of this major cost impact would be 

helpful. 

158
 
 

 

 

 

Response: INL’s new woody feedstock handling system was designed to deliver feedstock to 

the gasifier at a lower cost, by optimizing the sequence and type of operations. Since drying 

equipment in the new design was changed from the previously used rotary dryer, a direct 

comparison is not applicable. Some information is provided here to give an idea of the cost 

impacts. If the feedstock cost is changed from $61.57/dry ton (in this design) to $50.70/dry 

ton (as shown in a previous version of the MYPP [DOE MYPP November 2010]) the MESP 

will decrease by 6.4%. Adding the capital cost for a rotary dryer, which was included in the 

previous version of the model, increases the MESP by 5.5%. It can be said that the increased 

feedstock cost is nearly offset by the inclusion of the dryer in the delivered feedstock cost. 

The major benefit of the new handling system is the lower moisture content. The impact of 

this change is captured in the sensitivity analysis. 

5.	 Comment: Does the woody feedstock price include a realistic cost of storage? The high onstream factor demands significant on-site feedstock storage (for example, two-week storage 

may require as much as 75,000 m3). A contingency factor added to the theoretical INL 

feedstock cost, or at least an uncertainty range associated with it, would also be in order. 

Response: Harris Group Inc. has also included land requirements for feedstock handling and 

storage in estimating the plant footprint and land cost (refer to Appendix L). INL has 

included costs for receiving, storage, queuing, and preprocessing at the plant in Table 6 of the 

report. INL will soon release a report detailing the design of the woody feedstock handling 

system. 

6.	 Comment: The use of hot flue gases to dry down to such moisture levels in a rotary drum 

type dryer would probably result in unacceptable emissions levels. 

Response: The design basis for feedstock drying is no longer a rotary dryer and the 

feedstock moisture at the gasifier in this design is 10% (up from 5% in the previous design). 

A flue gas scrubbing system was also added to the design (refer to the response for comment 

#2). 

7.	 Comment: Why not use a train instead of trucking in the biomass feedstock with 136 

deliveries per day. There are many potential reasons for this including traffic, local 

emissions, overall efficiency and cost, etc. 

Response: This will be dictated by local logistics and will be a suggestion for future
 

consideration.
 

8.	 Comment: The statement “debarked, delimbed logs are fed into the chipper which uses a 2 

inch internal screen” implies the bark and limbs are not utilized as feedstock. Is this correct? 

Response: The bark and limbs are also utilized as feedstock. This information and additional 

details on the feedstock supply system are provided in Appendix F. 

Area A200: Gasification 

9.	 Comment: Should the fresh water used by ash wetting be replaced by using process 

condensate, like water from water scrubbing unit? 

159
 
 

 

 

 

 

 

Response: Concentrated brine waste from the reverse osmosis brine recovery system was 

utilized for ash wetting in the final design case. This final design change reduced the 

wastewater and water consumption rates as shown by the “Optimized Water Consumption 

Case” in the Sustainability Metrics section. 

10. Comment: Steam is indicated as a fluidization medium for the gasifier and tar reformer. This 

not only appears unnecessary, but is also counter-productive as described in a recent patent 

application. 

Response: NREL authors understand that the fluidization of gasifier solids in the absence of 

supplemental steam has been demonstrated. However, this process scheme was not used in 

the design case since there is not substantial experimental information in the public domain 

to predict the gasifier outputs for such a configuration. Furthermore, this report represents an 

update to the 2007 Phillips et al. report (Phillips et al. 2007) and major changes to the process 

in the front section of the plant were not considered. The primary focus of this design update 

is to present the modifications in the alcohol synthesis and tar reforming areas along with 

updated capital costs for the plant. NREL plans to look at the technoeconomics of dry 

gasification and reforming when more data becomes available. 

11. Comment: The olivine attrition basis described and make-up allowance for the riser type 

DFB (dual fluidized bed) gasifier appears to be too low. 

Response: The olivine make-up rate due to attrition and system losses is assumed to be 0.011 

wt % of the circulation rate (99.989% overall efficiency of two-stage cyclone systems). The 

per-stage efficiency required to achieve this overall efficiency is less than 99%. Single-stage 

cyclones in similar dual-fluidized bed designs, like fluid catalytic cracking units, typically 

achieve greater than 99.9% efficiency (Buell Refinery Cyclones 2008). 

12. Comment: Suggest elevating the pressure of the front end gasifier and tar reformer units. 

Response: The operating pressure was increased from 25 psia to 35 psia (at the gasifier inlet) 

per recommendation. The purpose of this change is to ensure adequate available pressure 

drop for the product vapor from the tar reformer to flow through the syngas cooling circuit 

and syngas scrubber to the suction of the compressor without risk of bottlenecking the 

gasifier and tar reformer systems. 

13. Comment: Are there size estimates for the gasifiers? What would be the new size of the 

gasifiers when doubled to 1,000 MT/day per train? 

Response: Sizing calculations for the dual fluidized-bed systems were not provided to NREL 

as part of Taylor Biomass Energy’s quotation for the packaged systems. However, Taylor 

Biomass Energy provided NREL with a basis for scaling the costs for the gasifier and tar 

reformer systems from the quoted 500 MTPD capacity to 1,000 MTPD. 

14. Comment: What is the basis for the steam-to-feed ratio and the olivine per wood ratios? 

Perhaps this is covered in other NREL reports (Spath, Mann, et al. or Philips et al.) but a 

reference if nothing else would be helpful. 

Response: The steam-to-feed ratio in the gasifier is within the range of experimental runs 

using the BCL indirect gasifier. The base case design provides enough steam for fluidization, 

160
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

but at the same time does not exceed a steam-to-carbon ratio of 2, which is adequate for a 

fluidized bed tar reformer. Adding more steam in the tar reformer would increase the H2-toCO ratio in the syngas, adversely affecting ethanol productivity (as predicted by The Dow 

Chemical Company’s alcohol synthesis model). The olivine-to-biomass ratio is dictated by 

the energy balance and thermal profile of the gasifier. Heat transfer takes place by moving 

hot solids from the char combustor to the gasifier. However, there are temperature limitations 

for the char combustor. The char combustor operates at 1,808°F and the gasifier operates at 

1,596°F in the base design case. The heat capacity of the circulating solids should be 

sufficient to maintain the desired gasifier temperature without exceeding the temperature 

limits in the char combustor. The values used in this design are similar to those presented in 

previous NREL reports mentioned in the comment. 

15. Comment: What is the basis for 2% LHV in gasifier heat loss? Is this a scaled result? If so, 

how? (Volume-to-area ratios?) 

Response: NREL authors did not perform detailed calculations for the heat loss. 2% is 

simply a reasonable estimate. Heat loss estimated for a 500 TPD indirect (BCL) gasifier was 

approximately 4% (Feldmann et al. 1988). With improvements in insulation and scale-up to 

1,000 MTPD, 2% heat loss is achievable. The impact of a higher heat loss of 4% is captured 

in the sensitivity analysis. 

Area A300: Gas Cleanup 

16. Comment: As the tar reformer is a significant research area, it might be helpful to have some 

discussion about the tar reformer – it seems like the costs aren’t substantially different from 

the separate tar cracker/steam reformer, but there must be a significant savings in energy 

from not cooling down raw syngas then heating it back up. 

Response: As mentioned in the comment, the fluidized tar reformer in this design converts 

tars, methane, and other hydrocarbons in a single operation. The capital cost for a single step 

fluidized bed reforming system will be less than or comparable to the cost of a design 

consisting of separate tar cracker and steam reformer systems, depending on the chosen 

configurations (fixed bed, fluid bed, etc.). A single step process is an elegant conversion 

method that prevents tar losses, does not require intermediate cooling and reheating, and 

minimizes associated capital costs and energy losses relative to two separate systems. NREL 

is further evaluating various options for tar reforming, internally and with industrial partners, 

and will conduct pilot plant tests in 2012. 

17. Comment: What is the design basis for sending 75% of the recycled purge syngas to tar 

reformer and the remaining to the reformer catalyst regenerator? Or what is the specific 

effect of this split on the methane content? If splitting more purge gas to reformer, will it 

increase the final products yield? Artifact of power neutrality? 

Response: Power neutrality requires the use of a portion of the syngas to provide heat via 

combustion. The design consideration here is to preferentially route unconverted syngas 

purge to the catalyst regenerator rather than raw syngas from the gasifier, as the purged 

syngas is sulfur free. This also prevents the buildup of chemical species that may not be 

161
 
 

 

 

 

 

 

 

 

 

converted in the tar reformer. These species are effectively removed from the plant by
 

combustion in the catalyst regenerator and purge through the flue gas system.
 

18. Comment: Space velocity indicated for reformer looks reasonable. 

Response: Thank you for confirming this. 

Area A400: Syngas Compression, Alcohol Synthesis, Acid-Gas Cleanup 

19. Comment: Suggest that you get The Dow Chemical Company to at least provide a budgetary 

quote for the DEPG acid-gas removal system as opposed to the approach you used. I'm 

concerned that costs maybe inadequate if one needs to be guaranteed by the licensor. 

Response: The feasibility of using the DEPG physical solvent at the process conditions 

specified was confirmed by The Dow Chemical Company and UOP LLC. A preliminary 

review of the design basis and cost estimate by The Dow Chemical Company indicated 

conservatism in NREL authors’ assumptions. The Dow Chemical Company suggested 

contacting UOP for a detailed assessment. NREL authors sent design information to UOP for 

evaluation but did not receive a response after multiple requests. 

20. Comment: The work of adding the recent catalyst information from The Dow Chemical 

Company is valued. It would be helpful to understand how this catalyst differs from the 

hypothetical catalyst described in the previous report. 

Response: NREL did not have access to the catalyst or kinetic data for the 2007 design. 

Assumptions on catalyst performance were based on data published in the literature. Since 

the 2007 design, NREL has entered into a Cooperative Research and Development 

Agreement (CRADA) with the Dow Chemical Company, allowing NREL limited use of The 

Dow Chemical Company’s kinetic model for alcohol synthesis. The kinetic model has 

predictive capabilities for the behavior of the catalyst, which is especially useful for 

predicting the impact of recycle streams and inert components in the highly integrated 

process. 

21. Comment: It seems that the H2S/CO2 recovery system is rather complex and costly. 

Consideration to commenting on this would be helpful. 

Response: The H2S/CO2 recovery system appears complex as it is designed to accomplish 

two major objectives. The first objective is to control the CO2 concentration in the total 

alcohol synthesis reactor feed. This is efficiently accomplished in the unconverted syngas 

recycle loop where the CO2 partial pressure is high after alcohols have been removed by 

condensation. The high CO2 partial pressure minimizes the DEPG circulation requirement, 

thus minimizing associated capital and operating costs. The second objective is to allow 

control of the H2S concentration in the alcohol synthesis reactor, required to maintain catalyst 

sites in a sulfided state. The amine acid gas enrichment and methanol absorption systems are 

designed to accomplish this objective. The acid gas enrichment unit takes the weak (low in 

H2S) acid gas stream from the DEPG system, rejects a significant amount of CO2, and 

generates a more highly concentrated H2S stream. This highly concentrated stream allows for 

H2S absorption into a pure methanol stream. After the absorption, the methanol and H2S are 

recycled to the alcohol synthesis reactor. 

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22. Comment: In the 2007 NREL design report, there is no syngas recycling in the mixed 

alcohol synthesis process and all the unreacted syngas is recycled to the tar reformer process. 

In this report, 88% of unreacted syngas is recycled to the synthesis reactor. What is the 

reason for this change? 

Response: This change resulted from the incorporation of the alcohol synthesis kinetic 

model from The Dow Chemical Company into the current design. Using this tool, NREL has 

observed that the most economically attractive operating conditions are high pressure and 

low per-pass conversion, which favor ethanol yield and selectivity. The low per-pass 

conversion requires a high recycle ratio. Recycling large volumes of unconverted material to 

the tar reformer requires significant decompression and recompression of the syngas, which 

would significantly increase electrical consumption and negatively impact the overall product 

yield. By keeping the majority of the unreacted gases in the synthesis loop, recompression 

requirements are relatively low, improving the process economics. 

23. Comment: Methanol absorption of H2S and recycle to alcohol synthesis is used to maintain 

the H2S partial pressure for the catalyst. Why not just use the DEPG process to control the 

removal of H2S and thus maintain the H2S in recycle gas and the H2S partial pressure? 

Response: The solubility ratios for CO2 and H2S in DEPG do not allow this simplification. 

The primary objective for the DEPG system is to control the concentration of CO2 to the 

alcohol synthesis reactor. The high solubility of H2S in DEPG results in near complete 

removal of the species from the recycle stream to the synthesis reactor. If sulfur content in 

the raw syngas (from raw biomass feed) is not adequate to maintain the minimum partial 

pressure of H2S for sulfided catalysts, then another H2S source must be maintained. NREL 

was opposed to adding an external sulfur source such as DMDS (dimethyl disulfide) for H2S 

maintenance as this would negatively impact the sustainability of the process. An internal 

H2S recycle loop was established in this design by first generating a high-purity H2S stream 

(through amine-based acid gas enrichment) and then recycling the material by contacting the 

H2S-rich stream with methanol. 

24. Comment: Is the CO2 mole fraction of 13.5% in the inlet syngas too high? Would higher 

CO2 removal in the DEPG process lower the capital costs of the syngas compressor and 

mixed alcohol synthesis, or perhaps this is offset by increased DEPG costs. 

Response: There is potential for further optimization by adjusting the CO2 concentration. 

NREL has observed through many model predictions that varying the CO2 content can have 

an economic impact of a few cents on the MESP. Based on NREL’s observations, the 

optimal CO2 concentration lies between 10% and 13.5% (base design concentration). Several 

factors impact the optimal operating point: 

•	 The capital and operating costs associated with the DEPG system increase with 

decreasing CO2 partial pressure. 

•	 The energy penalty from recompression of CO2 recycle is marginal because the 

recycle compressor has to compensate only for the pressure drop through the reactor 

and associated heat exchangers. 

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•	 Removing CO2 results in increased partial pressures of other reactant species, which 

helps improve yields in general. On the other hand, because CO2 participates in the 

reverse water gas shift reaction (CO + H2O  CO2 + H2) in the reactor, higher CO2 in 

the recycle loop helps to maintain higher CO in the reactor feed, improving ethanol 

productivity and selectivity. 

Technical targets are set for catalyst performance improvements by 2012 (refer to Appendix 

I), which are incorporated into the base case design. Adjusting the CO2 concentration is one 

of the variables used by NREL to tune model predictions to align with technical targets. As 

mentioned, NREL’s model runs indicate that the base case CO2 concentration is near the 

optimal operating point. NREL will continue to optimize the model in future State of 

Technology reports if opportunities for reducing MESP are recognized. 

25. Comment: Not sure why DEPG is used for CO2 removal instead of an amine system. 

Normally DEPG is better for high pressure gases and an optimization evaluation would be 

necessary when considering compression of the gas. 

Response: The decision was made to utilize a physical solvent for acid gas removal when the 

alcohol synthesis pressure was increased to 3,000 psia (to optimize product yields). Also 

refer to responses to comments #21 and #24. 

26. Comment: The report says “The perturbed-chain statistical associating fluid theory (PCSAFT) package was used to model physical solvent absorption processes in a separate 

simulation.” This is an odd choice for the DEPG. What is the basis? 

Response: The starting point of NREL’s DEPG model was an Aspen Plus example that uses 

the PC-SAFT property method and already has physical property parameters available. 

27. Comment: The location of the DEPG acid gas system is odd in that it is only on the residual 

syngas recycle, not on the fresh feed to the synthesis reactor. 

Response: Complete H2S removal is not required upstream of the alcohol synthesis reactor. 

The primary objective of the acid gas removal system is to control the CO2 concentration at 

the throat of the synthesis reactor within a range that does not negatively impact conversion 

and selectivity to ethanol. As discussed in the response to comment #24, it is also beneficial 

to feed some CO2 to the reactor to prevent loss of CO through the water gas shift reaction, 

which is very fast on metal sulfide catalysts. As discussed in the response to comment #21, 

the syngas recycle loop is currently believed to be the optimal location for the DEPG system. 

28. Comment: I noticed you are using an absorption refrigeration system. Was there any trade 

study done on vapor compression refrigeration (whose COPs are much higher) and perhaps 

drive it with the onsite steam Rankine cycle (or use a higher temperature packaged organic 

Rankine cycle power generator)? 

Response: NREL incorporated a vapor compression system in earlier design cases. The high 

power requirements of the vapor compression system and availability of heat streams in the 

plant prompted NREL authors to explore absorption refrigeration. The advantages of the 

absorption refrigeration system fit well and resulted in favorable economics compared to the 

vapor compression system scenario. Note that the plant design is energy neutral, meaning no 

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electricity is imported or exported. Therefore, energy benefits directly result in increased 

product yields and decreased MESP. The design includes a steam Rankine cycle to produce 

electricity. NREL did not consider an organic Rankine cycle for this design. 

29. Comment: Compressor efficiency: In the 1st paragraph, polytropic compressor efficiencies 

are given. What is the total efficiency (including mechanical/motor efficiencies) employed? 

Response: The six compressor stages have different efficiencies that were obtained from 

vendor quotes. A 98% mechanical efficiency was applied on top of the polytropic 

efficiencies quoted. A weighted average efficiency per stage (weighted by power required 

per stage) is approximately 82%. 

30. Comment: It is stated that alcohol synthesis is highly exothermic. Please place a ΔH value 

for a representative synthesis reaction or a range for Eq. 4. 

Response: For the ethanol synthesis reaction represented by: 

2CO + 4H2  C2H5OH + H2O ΔH = -61.2 kcal/mol 

This value was calculated using Aspen Plus with references of 25°F, 1 atm, and vapor phase. 

31. Comment: More details on synthesis reactor sizing and steam-side heat transfer coefficients 

would be helpful. 

Response: The reactor was sized based on commercially available tube dimensions and 

maximum tube length of 60 feet. Through an optimization exercise, NREL obsverved that 

two parallel reactors provided the optimal balance between capital costs and reactor 

performance. The maximum number of tubes per reactor is 5,000 based on fabrication 

constraints. The space velocity and reactor conditions (such as shell side temperature) were 

manipulated in the integrated simulation until optimal conditions were achieved (within the 

limitations described above). The results shown for the base case reflect this methodology for 

reactor sizing. The steam side heat transfer coefficient was conservatively assumed to be 500 

W/m2/K (88 Btu/h/ft2/°F). 

32. Comment: It is stated that the inlet pressure to the first expander had to be dropped from 

2,900 to 1,500 psi. This represents potentially large exergy destruction. I believe this loss 

could be mitigated by looking for places where a gas ejector (aka jet pump for liquid-phase 

streams) could help with compression, such as at the inlet to the intercooled syngas 

compressor train. This could potentially lower the plant parasitic power. Please feel free to 

contact me regarding this if you have not already considered it. 

Response: The large pressure drop (2,900 to 1,500 psi) is only in the stream that is purged 

from the recycle loop. This purge stream cannot be mixed with the fresh syngas feed, which 

is compressed after the quench system. 

33. Comment: Expander efficiency was 75%...is this isentropic, total efficiency, other? 

Response: The isentropic efficiency is 75% and the mechanical efficiency 97%, resulting in 

overall efficiency of 72.8%. The same is assumed for the steam turbines. 

34. Comment: What is the design basis for purging 12% unreacted syngas? If the built-up of 

inert gas in recycle loop is the only basis, what is the limit of the inert gas built up? If 

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purging more and recycling it to the tar reforming, what are the effects of it on capital and 

operating cost? Is this an artifact of power neutrality? 

Response: The purpose of the unconverted syngas purge is to control the build-up of inert 

components that are not removed effectively by the DEPG acid gas removal system. The 

presence of inert compounds in the alcohol synthesis reactor impacts the partial pressures of 

the reactants. Significant reductions in reactant partial pressures will have a negative impact 

on yields as predicted by The Dow Chemical Company’s kinetic model. The basis for the 

unconverted syngas purge is the simultaneous optimization of the reactor model yields, 

energy requirements for recompression, and capital costs. 

35. Comment: For the statement “Light hydrocarbons, methyl esters, and aldehydes are 

produced in smaller quantities through similar chemical routes,” do you include those 

compounds in the model? 

Response: The model includes small amounts of those species. However, the quantities are 

not significant and are not included in the material streams shown in Appendix E. 

Area A500: Alcohol Separation 

36. Comment: It is mentioned that the condensed methanol/water is recycled to the gasifier. But 

in Appendix E (process flow diagrams), it shows that this stream is recycled to the MAS 

synthesis preheater. Which one is correct? 

Response: The condensed methanol/water stream from the molecular sieve system (stream 

513L) is recycled to the gasifier. The ambiguity has been resolved in the final report. 

Area A600: Steam System and Power Generation 

37. Comment: While I support the vision of not utilizing natural gas and/or external power, I did 

not see any capital allocated to start-up the facility under this scenario. In specific, the 

gasifiers and associated equipment will be refractory lined, and will thus require gradual 

heat-up during start-up and after repairs. This is traditionally provided by natural gas, yet I 

saw no allocation of capital towards bulk storage or distribution of natural gas and/or LPG. 

Further, since internal power generation is based on the combustion of compressed synthesis 

gas and subsequent steam generation, the feed systems and gasifier will require external 

power for start-up. The capital estimate did not appear to include provisions for the estimated 

5-15 MWh of electricity (depending on whether the first stage of compression is necessary to 

power the system). 

Response: Equipment for start-up of the gasifier trains is included in the scope of the 

quotation from Taylor Biomass Energy. In addition, NREL has included approximately 

$10MM in direct costs ($15MM after indirect cost factors are applied) to cover site 

development costs. It is assumed that electrical supply from the grid to the plant will be 

covered under site development. Transformers, switching stations, and other power system 

infrastructure are assumed to be fully covered in equipment installation factors as 

summarized by Harris Group Inc. in Appendix K. 

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Area A700: Cooling Water & Other Utilities 

38. Comment: The study assumes power self-sufficiency and a modest power export. This 

assumption needs to be supported by the process concept by including costs for power 

transforming stations and power switching stations in the cost estimate. Other OSBL 

equipment, such as flare systems, waste storage, etc., also should to be included. These 

additional offsites will probably be a sizeable fraction of the assumed Area 700 cost, which is 

currently a diminutive 1.3% of the TIC. I understand that some supporting equipment and 

infrastructure is already included in the assumed installation factors, utility and waste 

processing costs, as well as Area 700 components. However, as comparison, NETL work 

cites OSBL costs totaling between 11% (NETL 2007) and 21% (NETL 2005) of the TIC for 

IGCC plants. The large difference in OSBL cost contribution between this work and existing 

literature requires clarification. 

Response: NREL authors believe the capital costs presented in the report adequately cover 

allowances for OSBL facilities based on the following: 

•	 The NETL IGCC OSBL costs referenced above (as percentages of total installed cost) 

include steam, power, condensate, and wastewater systems. When costs for these 

systems are considered for the thermochemical ethanol design (Areas 600 and 700), 

the installed cost relative to TIC is approximately 19% ($55.7MM / $296.7MM), 

which is at the high end of the NETL range referenced above. 

•	 Transformers, switching stations, and other power system infrastructure are assumed 

to be fully covered in equipment installation factors as summarized by Harris Group 

Inc. in Appendix K. 

•	 An emergency flare system is included in costs of the gasifier and tar reformer units 

as quoted by Taylor Biomass Energy. 

•	 NREL has included approximately $10MM in direct costs ($15MM after indirect cost 

factors are applied) to the estimate to cover site development costs such as roads, 

logistical access points, wastewater treatment facilities, and other requirements for a 

green-field site. 

39. Comment: The assumption that an existing publicly owned wastewater treatment system 

(POTW) will be available for the majority of the nth facilities is overoptimistic. In fact, the 

overall percentage of capital allocated towards environmental controls is well below industry 

averages. Based on my project background, one-third of the capital will typically be 

necessary to meet environmental regulations. Examples of potential issues with this project 

include raw water treatment, vent emission controls, waste water treatment, storm water, and 

spill control. 

Response: NREL has incorporated a flue gas scrubber for control of flue gas emissions and a 

more rigorous water treatment and recycle system into the design after the December 2010 

review draft. In addition, NREL has included approximately $10MM in direct costs ($15MM 

after indirect cost factors are applied) to cover site development costs. It is assumed that 

wastewater storage and treatment facilities will be covered under site development. 

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40. Comment: What are the sources for “water” in the Energy Balance section? 

Response: The energy analysis table in the review draft has been replaced by a pie-chart in 

the final version (Figure 14). The major sources of water contributing to the energy analysis 

chart (LHV basis) are makeup to cooling tower and boiler system. The source of this water 

will be site specific. 

41. Comment: Where are the wastewater treatment costs in the off-site facility from gas 

scrubbing accounted for in the economics? How is this transported (piped, trucked, train)? 

Response: Additional equipment (e.g., reverse osmosis) was included in order to minimize 

quantities sent to wastewater treatment. Per reviewer comments, NREL has increased the 

site-development costs in the estimate to cover items such as an on-site wastewater treatment 

plant. Operating costs associated with the on-site wastewater treatment plant are also 

included in the analysis. 

General 

42. Comment: The 96% availability is extremely high given that current NGCC and liquid-fed 

IGCC plants (much lower complexity) strive to achieve 90% availability with a single-train 

configuration. A 2-week turnaround leaves no room for unplanned down time. This high 

availability needs to be supported by additional investment in spare trains, intermediate 

storage, automation, redundant power systems, etc. A critical equipment analysis should be 

conducted to identify equipment that needs to be fitted with redundant trains, or 

intermediates that need to be stored, to ensure uninterrupted operation. The sparing of the 

gasifier, reformer, and synthesis reactor is commended. However, I suggest at least a 

discussion of whether spare trains for the syngas train (i.e., compressor, gas scrubbing, acid 

gas removal) were included, and if not, the supporting rationale (e.g., high cost of 

compression, existing guard beds, etc.). 

Response: The assumption that an nth thermochemical conversion plant will achieve 96% 

utilization is based on two factors. The first is simple – consistency of assumptions between 

the biochemical and thermochemical platform analyses. However, NREL recognizes the 

point that thermochemical conversion is a more severe operation than biochemical 

conversion, which necessitates additional support for the assumption. Petroleum refinery 

operations like fluid catalytic cracking and fluid coking are similar to the biomass 

gasification and tar reformer processes in that they are high temperature (although lower than 

the gasification reactor), dual fluidized bed systems. Refiners have gained tremendous 

improvement in operating reliability of these refinery applications as the technologies have 

matured. Similar reliability improvements can be expected for biomass conversion processes 

as the technology matures. John S. Magee states, “It is not uncommon today to have FCC 

units complete three-to-four-year runs between major turnarounds with nearly 100 percent 

stream factor” (Magee 1993). The on-stream factor assumed for the thermochemical design 

allows a 30-day turnaround per two-year run. NREL has provided MESP values below 

corresponding to varying on-stream factors. A 2% decrease in assumed on-stream factor 

increases the MESP by roughly 3 cents per gallon. 

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Assumed On-Stream Factor MESP ($/Gallon Ethanol) 

96 % (Base Design Case) $2.05 

94% $2.08 

92% $2.11 

43. Comment: The final area of concern with the project was with the estimated on-stream time. 

In my experience, the assumption of 96% on-stream reliability (without any redundant 

systems) would be unrealistic. Further, the best numbers I have seen to-date on biomass 

gasifiers is 92% after three years of continuous operation. And while I can appreciate that 

there is room for improvement above 92% for an nth facility, 96% is too high, especially 

given the planned maintenance budget and operations staffing. In my experience, to achieve 

96% on-stream reliability requires redundant critical systems and a well developed 

preventative maintenance program. 

Response: Please refer to the response to comment #42. The maintenance budget was also 

increased (see comment #54). 

44. Comment: When you do the integration, keep in mind the cost of pipe and ductwork. You 

don’t want to be routing pipe and duct all over the plant. 

Response: NREL considered piping costs when incorporating the exchanger network into 

the plant design. During the design of the heat exchanger network some of the large 

exchangers were designed to exchange heat within the same plant area. For example, most of 

the alcohol synthesis inlet stream preheating and product stream cooling were done by 

exchanging heat between those streams. However, it was not possible to limit all the process 

heat exchange within the same unit operation area. For example, the hot syngas and flue gas 

streams have large amounts of heat available in a wide temperature range. Efficient recovery 

of that heat is possible only by considering the heating needs of the entire plant. The design 

of the exchanger network will be an area of focus during detailed design of commercial 

thermochemical plants in order to maximize efficiency while minimizing cost. 

45. Comment: Though you need it for the DCF (discounted cash flow) analysis, I would not 

include working capital in the total project cost. 

Response: In the final version of the report, the sum of the fixed capital investment (FCI) 

and working capital yields the total capital investment (TCI). This nomenclature was used for 

consistency with NREL’s biochemical ethanol design report. 

46. Comment: Look at replacing the water usage versus air coolers. Need to look carefully at the 

tradeoff between capital and operating costs here. 

Response: NREL has incorporated several air coolers in the design to reduce water
 

consumption for plant sustainability. 

47. Comment: Please provide basis for converting to 2007 dollars. 

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Response: The conversion to 2007 U.S. dollars is accomplished through Equation 2 and the 

cost index values in Figure 2. For example, if equipment costs were quoted at $1MM in 

2005, the costs would be adjusted to $1.12MM for the 2007 basis ($1MM * 525.4 / 468.2), 

where 525.4 and 468.2 are the cost index values for 2007 and 2005, respectively. The basis 

for converting operating costs is noted in Table 29 (footnote). Labor costs are indexed (if 

necessary) to values from the U.S. Bureau of Labor Statistics (http://data.bls.gov/cgibin/srgate CEU3232500008). 

48. Comment: Identify the basis for indirect costs and project contingency of 60%. 

Response: The indirect cost factors are consistent with those applied in the biochemical 

ethanol design report. The values were recommended and agreed upon through the peer 

review process. For this design, 60% of total direct costs (TDC) correspond to 37.5% of fixed 

capital investment (FCI), which is consistent within the 16% to 61% of FCI range presented 

by Peters, Timmerhaus, and West (2004). 

49. Comment: It may be worth looking into the economic interplay of the energy import/export 

and alcohol production with the ultimate goal to derive a globally optimum design that 

minimizes total production cost. This, however, may not have been DOE’s directive in regard 

to this work. 

Response: The design consideration that biorefinery facilities will be energy self-sufficient 

for steady state operation is consistent between biochemical and thermochemical ethanol 

pathways. Though economics may improve by purchasing electricity and natural gas for 

plant operations, the long-term sustainability of the processes will suffer. The consistent 

assumption allows for easy comparison (a level playing field) among the various technology 

platforms without adjustments. 

50. Comment: The study yields a lower Lang factor (3.4) than those universally accepted for 

any “nth plant” study. For example, Peters, Timmerhaus, and West (2004) recommends a 

TCI-based Lang factor range between 4.7 for solids processing plants and 6 for fluids 

processing. Along these lines, a discussion explaining the lower than expected Lang factor is 

warranted. 

Response: There was a typographical error in the draft version of the design report. The 3.4 

value is not the Lang factor as defined in Peters, Timmerhaus, and West. The value 

represents the quotient of the total capital investment (TCI) divided by the total purchased 

equipment costs (TPEC) including the OSBL equipment (Areas 600 and 700). Lang factors 

for FCI and TCI (defined below) have been included in the revised report. 

Design 

Report 

Values from Peters, 

Timmerhaus, and West 

Solid Solid-Fluid Fluid 

FCI Lang Factor = FCI / ISBL TPEC 4.12 4.0 4.3 5.0 

TCI Lang Factor = TCI / ISBL TPEC 4.32 4.7 5.0 6.0 

51. Comment: Cost escalation is typically considered for a project that may take 1-2 years to 

complete engineering work on and 2-4 years to build. Thus, a 3%/yr escalation factor—in 

excess of normal inflation—may result in ~10% in additional investment costs. A short 

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discussion explaining whether cost escalation was considered and, if not, the reasons for not 

including it may prove useful. 

Response: NREL did not consider cost escalation (or profit escalation) from inflation as all 

monetary values are presented on a constant basis of a cost year (2007 U.S. dollars in this 

case). The MESP is calculated by solving for net present value at start-up (after construction) 

equal to zero. 

52. Comment: KWh-to-hp conversions appear to be at 95% efficiency versus the stated 75% 

efficiency (e.g., 19,683 hp generated to produce 13,953 KWh of electricity). 

Response: In the review draft version this was left over from the previous version of the 

design, in which several pieces of rotating equipment were steam driven. In the current 

design all rotating equipment is electric motor driven. The efficiencies applied in the model 

are consistent with the values listed in this report. 

53. Comment: It is stated in Section 2 that the preferred feedstock for these facilities would be 

wood biomass. Given the sustained demand for wood, and the required capture radius, it 

would seem reasonable to assume that the majority of the 350 theoretical facilities would be 

located within rural/remote areas. I believe that the expected costs for land and site 

development are underestimated. Specific gaps would include necessary upgrades to the 

transportation infrastructure around the facility, as well as those necessary on the site itself. 

This would ultimately include elements such as roads, bridges, rail systems, wetlands 

mitigation, wells, and outfall systems. As such, I believe that the costs for developing a 

green-field site in a remote area will exceed 6% of the TPEC, or in this case $8.5MM. 

Response: Since the review draft, NREL has included approximately $10MM in direct costs 

($15MM after indirect cost factors are applied) to cover site development costs in addition to 

the land cost estimate of $1.6MM based on the plant footprint estimate by Harris Group Inc. 

(refer to Appendix L). 

54. Comment: Given a replacement asset base of $470MM, the projected 2% of RAB as a 

maintenance budget will be insufficient to maintain the facility in good repair (e.g., the entire 

biomass through syngas generation system is high maintenance). Budgets notwithstanding, I 

also do not believe that two technicians per shift (or worse yet, eight working M-F on day 

shift) will be sufficient to maintain the facility at the stated operating factor. 

Response: Based on this comment, NREL increased the maintenance budget by 50% to 3% 

of fixed capital investment (FCI) annually and doubled the number of maintenance 

technicians from 8 to 16 (2 per shift to 4 per shift). 

55. Comment: The couple of things I did notice were on the financial front. First, even for the 

nth plant, I think 60% leverage is very high. That is the kind of leverage we saw in 1st gen 

and unless the plants have hedged off take I think that will be difficult. I would be thinking 

more like 40%–50% leverage. Also, even for the nth plant, I think the cost of debt and equity 

are going to be low (unless there are off take agreements). Equity would likely seek a 12%– 

15% return and I would look at debt in the 8%–10% range. 

Response: Although this is a valuable comment, NREL maintained consistency of financial 

assumptions with similar reports produced for biomass conversion under the DOE Office of 

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the Biomass Program’s umbrella. Sensitivity scenarios based on these comments on financial 

assumptions are presented in Section 5 of the report. A sensitivity scenario is provided below 

to address the specific recommendations of the reviewer around project financing. With the 

equity increased to 55% of total financing, return to equity investors increased to 12%, and 

interest rate on debt increased to 9%, the resulting MESP is $2.25 per gallon of ethanol. 

56. Comment: Was the plant size scale-up based on scaling the total installed costs or did you 

take into account pieces that need to be number up (gasifier, tar reformer). A note to that 

effect in the sensitivity section would be helpful. 

Response: The scale-up sensitivity case scales each piece of equipment or packaged unit 

based on the associated scaling factor. For the gasifier and tar reformer trains, the maximum 

capacity per train is limited to 1,000 metric tonnes per day. The 10,000 metric tonnes per day 

sensitivity case specifies 10 gasifier and tar reformer trains in the equipment list. 

57. Comment: Accuracy of estimate. I've sent the AACE document separately that I've used as a 

basis to say that the accuracy is really within 50%, not 30%. 

Response: NREL authors feel that with the use of recent experimental data and vendor 

quotes the uncertainty in the capital cost estimates is significantly lower than 50%. However, 

even if the uncertainty in the total capital cost is assumed to be the high value of +50%, the 

corresponding uncertainty in the MESP is approximately +30%. NREL authors mention 

±30% as the range of accuracy of the MESP (Appendix I), which corresponds to the 50% 

uncertainty in capital cost estimates. The sensitivity analysis in Section 5, is based on the 

expected accuracy values presented in Peters, Timmerhaus, and West (2004). 

58. Comment: The assumption of “nth plant” for a processing scheme that has yet to see its first 

commercial application as a whole makes any numerical prediction of estimate uncertainty 

questionable. The statistics connecting a feasibility or budgeting estimate to their expected 

accuracy ranges only apply for projects similar to those that statistics were built upon (i.e., 

commercially demonstrated technology). For example, the 2007 NREL Mixed Alcohol 

Report—an R&D target study—predicts $1.01/gal breakeven product value with $137MM 

TIC and a (+30)–(-10)% estimate accuracy, yet this year’s re-evaluation produces costs that 

are twice as large, well outside the previously assumed uncertainty range. This increase in 

estimated costs is a normal phenomenon as pre-project technical development progresses. 

Moreover, the practical feasibility of research targets is inherently uncertain making a 

prediction of accuracy difficult at this stage. Along these lines, I suggest qualifying the 

estimate class as a “Class 4–type (feasibility/study)” or “Class 3–type (budget authorization)” 

based on guidelines developed by the Association for the Advancement of Cost Engineering, 

but without including a numerical accuracy range. 

Response: Please see the responses to related comments #57 and #61. Also note that the 

$1.01/gal projection was based on a lower feedstock cost ($35/dry ton) and is in 2005 U.S. 

dollars. The MESP was since revised to $1.57/gallon in 2007 U.S. dollars, using a higher 

feedstock cost of $50.70/dry ton, as documented in Table B-5 of the November 2010 MultiYear Program Plan (DOE MYPP November 2010). The current MESP of $2.05/gallon is 

roughly 30% higher than the previously projected $1.57/gallon. The authors are confident 

that because estimates in this report are based on experimental results and recent vendor 

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estimates, the results presented in this report are representative of nth plant costs for the 

technologies proposed in this design. It should be noted that there were significant 

uncertainties in the alcohol synthesis section of the 2007 Phillips et al. (Phillips et al. 2007) 

report. In 2007, the performance of the alcohol synthesis reactor was inferred and 

improvements extrapolated to 2012 based on literature information because that was the best 

available information at that time. In the current design, which is based on experimental 

results, the cost estimates are based on current understanding of the state of technology. 

Increases in capital estimates can be categorized as follows: 

•	 The increased capital estimates are primarily a result of obtaining recent vendor 

quotes through Harris Group Inc. For example, the recent estimates of gasifier and 

reformer installed costs are three times higher than those presented in the 2007 report. 

There are reviewer comments suggesting that the gasifier/reformer estimates may be 

high for nth-plant costs because they are based in today’s quotes and costs should 

decrease with maturity associated with the nth-plant (comment #62). There remains 

some scope for a downward revision of the gasifier/reformer costs. 

•	 The alcohol synthesis process conditions and flows were redesigned based on 

experimental and modeled data. The operating pressure is higher and conversions are 

lower than in the 2007 assumptions. The modified process led to higher capital costs. 

NREL was able to offset some of the increased capital costs in this area by improving 

the overall yields through improved process design and energy integration. There is a 

potential for the reduction of costs in this area with improvements in catalysis. 

•	 The detailed vendor estimates for the compressors also showed a marked increase 

from the 2007 estimates (partly due to the higher operating pressure). 

Although it is usually the case that cost revisions follow an upward trend, NREL authors feel 

that the projected nth-plant costs in this report are conservative in some areas and there is 

potential for a downward revision in the future. 

59. Comment: The only consideration I have a slight issue with is that of the “nth plant” design 

in that it helps in providing justification and support for early technology adopters and 

pioneer plants. I believe even these types of investors will have to be reassured that it is not a 

matter of 3 to 5 years down the road before a lower cost plant could be designed and put on 

stream. In my experience the “nth plant” will have benefitted from the knowledge-basis of 

previously designed plants resulting in a plant capital cost that typically falls on the lower 

portion of the cost vs. number of plants built “curve.” To this end, I do not think in its present 

state this design falls in the “nth plant” category. However, as I said it is not a major issue and 

given the thoroughness of your sensitivity analysis the concern with the “nth plant” 

assumption is most likely a moot point. 

Response: NREL authors agree with this observation and believe the estimates presented in 

this design may be on the conservative side of the range for nth plant costs. 

60. Comment: I think you may have over-valued the co-products. I can't tell the petroleumderived equivalent fuel price on a $/gal basis from the report. I would think that you would 

want to be equivalent to a residual fuel oil, which is normally valued less than crude oil (on a 

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$/gal basis) and about 60%–70% of gasoline. The $1.90 per gal is almost the same as the 

MESP of your major product. Since this co-product contribution is about 10% of the MESP it 

should be calculated with care. 

Response: Based on this comment, NREL authors reduced the co-product credit to 

approximately 90% of the calculated energy value. Refer to Table 32 for details on coproduct pricing. 

61. Comment: Regarding the comparison of market studies in Table 2: Were the TE model 

results independent of the market studies? Text below Table 2 seems to imply they were not, 

in which case the usefulness of the comparison is strongly limited for the present study since 

no biorefineries of this scale exist. 

Response: The paragraph following Table 2 discusses this issue, stating that there are more 

uncertainties (less specific information) in the study of pre-commercial technologies. Based 

on the comment, NREL authors added an additional sentence stating that there will be more 

uncertainties in the cost predictions from such studies. 

62. Comment: Using first-of-its-kind cost estimates for gasifier, tar reformer, and syn. reactor 

seems inconsistent given the nth generation plant TE model philosophy. While the sensitivity 

analysis results may capture some of this variability, economies of production (no. units per 

year) combined with the hardware production learning curve are likely to result in more 

significant cost reductions than the range explored in the sensitivity analysis section. 

Response: NREL authors agree that there is a possibility of further cost reductions because 

the initial basis for the gasifier/reformer cost estimate is current technology costs. It should 

be noted that NREL’s design configuration has slight variations from the Taylor Biomass 

Energy system in the reforming section, and it was assumed that some of the savings in the 

nth-plant will be absorbed by these variations in NREL’s design. 

63. Comment: In this section and periodically throughout the report, hardware performance 

assumptions or improvements over the 2007 report are made but justification is not given. It 

would be beneficial to understand NREL's reasoning behind performance improvements, 

such as those given in Table 10 and in Sec. 3.4.2. References and rationale behind such 

improvements should be presented. 

Response: Table 10 referred above shows the performance of the tar reforming catalyst. 

Significant progress was made in this area through targeted research since the previous report 

in 2007. The performance improvements shown are based on experimental results. Sec. 3.4.2 

refers to the performance of the alcohol synthesis catalyst, which is now based on 

experimental and modeled results for The Dow Chemical Company’s catalyst. The previous 

values were based on information from literature. The updated values shown in this report 

represent the current state of technology and improvements achievable by 2012. 

64. Comment: It would be very useful to understand the current technology status of hardware 

and have it accompany any rationale given for performance improvement expectations. For 

example, the use of Technology Readiness Levels (such as those used by Department of 

Defense, NASA, numerous aerospace corporations, and increasingly DOE–albeit informally) 

could inform one of the current status. 

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Response: This is a good suggestion and will be considered in future reports. The majority of 

the equipment and technology used in this design is commercial. The key areas that require 

further development are (1) reforming catalyst, (2) extended testing of the catalyst in a 

fluidized bed and detailed design based on the test results, (3) alcohol synthesis catalyst, and 

(4) scale-up of the gasifier from 500 MTPD to 1,000 MTPD. NREL and its partners will 

work on items (1), (2), and (3) for 2012. Commercial vendors will need to work on scale-up 

of the gasifier from 500 MTPD and a larger-than-pilot-scale demonstration of the reforming 

catalyst. 

65. Comment: What kinds of pinch temperatures are employed in the heat exchangers? 

Response: Details for all exchangers are included in Appendix A, Appendix B, and 

Appendix M.
 

66. Comment: Pinch Analysis section (of review draft). Perhaps 'temperature versus enthalpy' 

nomenclature is use in some circles, but I believe such figures are typically named T-Q 

diagrams, where Q is the amount of heat transferred on the x-axis. 

Response: The text has been modified to state “heat flow” instead of “enthalpy.” 

67. Comment: Energy Balance section (of review draft): This is slightly confusing. It would 

make sense to report an overall plant efficiency on an HHV-basis in addition to LHV-basis 

and while keeping your LHV energy balance regarding the intermediate process streams. 

Further, the statement, “The enthalpies calculated and reported...” is also a little confusing. 

Perhaps you mean the total enthalpy (hf + dh) as the heat of formation only includes the 

energy of formation for a compound and the latent energy (at the standard state). 

Response: NREL authors did not perform a detailed review of this section prior to the 

release of the December 2010 review draft. The text in this section has since been modified. 

As you point out, Aspen Plus does report total enthalpy. During review, NREL authors 

simplified the text because this section was unnecessarily wordy and trying to explain things 

readily available in textbooks. Energy balance is still reported on an LHV basis. NREL 

authors feel this is more valuable for capturing the negative impacts of water input into the 

process. Also LHV is an important quantity for end use as a transportation fuel (for internal 

combustion). NREL authors agree it will be useful to add another set of data on an HHV 

basis and will consider this in future reports. 

68. Comment: It should be noted in the text that the sensitivity to plant size assumes use of the 

same technologies as those of the larger plants. This is not necessarily a good assumption 

when scaling costs. Perhaps this is reasonable for a 600-MTPD biorefinery, but it could be 

debatable as, for instance, multiple dolomite beds and catalytic candle filters may be 

preferred over water-scrubbing or warm gas desulfurization in conjunction with water-gas 

shift membrane reactors over Selexol type of gas cleanup. This also depends on your 

assessment of the technology readiness of alternative technologies. 

Response: This is a very valid point. NREL authors added the note to the text. 

69. Comment: It would be useful to summarize the uncertainty of key plant process areas shown 

in the cost summary (Appendix C). For instance, the uncertainty associated with alcohol 

synthesis reactor shown in the sensitivity analysis tornado chart is in the neighborhood of 

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-14.4% and +15.5% on MESP (when summing all associated capital and catalyst costs). This 

could also be done for the feedstock and other areas. 

Response: The NREL authors felt that individual sensitivities shown in Section 5 provide the 

necessary information to get an idea of the impacts of uncertainties. A more detailed analysis 

of uncertainty ranges will be conducted in the future. A simple sum of uncertainties from the 

tornado chart for each process area will be misleading. 

Appendix O References 

Buell Refinery Cyclones. (2008). “Fluid Catalytic Cracking Cyclone Design and 

Modifications.” Buell Refinery Cyclones website. Lebanon, PA: Fisher-Klosterman, Inc., a 

CECO Environmental Company. http://www.buellrefinery.com/refining_details.htm. 

DOE MYPP. (November 2010). Biomass Multi-Year Program Plan. Washington, DC: U.S. 

Department of Energy. 

http://www1.eere.energy.gov/biomass/pdfs/biomass_mypp_november2010.pdf. 

Feldmann, H.F.; Paisley, M.A.; Appelbaum, H.R.; Taylor D.R. (1998). Conversion of Forest 

Residues to a Methane-Rich Gas in a High-Throughput Gasifier. Report DE88013138. 

Columbus, OH: Battelle. 

Magee, J.S. (1993). Fluid Catalytic Cracking: Science and Technology. Amsterdam, the 

Netherlands: Elsevier. 

NETL. (2005). Task 3 Gasification Plant Cost and Performance Optimization: Final Report. 

Work performed by Nexant under DOE Contract No. DE-AC26-99FT40342. Pittsburgh, PA: 

National Energy Technology Laboratory. 

NETL. (2007). Baseline Technical and Economic Assessment of a Commercial Scale 

Fischer-Tropsch Liquids Facility. DOE/NETL Report No. 2007/1260. Pittsburgh, PA: 

National Energy Technology Laboratory. 

Peters, M.S.; Timmerhaus, K.D.; West, R. (2004). Plant Design and Economics for Chemical 

Engineers, 5th Edition, International Edition. New York: McGraw‐Hill Inc. 

Phillips, S.; Aden, A.; Jechura, J.; Dayton, D.; Eggeman, T. (2007). Thermochemical Ethanol 

via Indirect Gasification and Mixed Alcohols Synthesis of Lignocellulosic Biomass. Report 

No. NREL/TP-510-41168. Golden, CO: National Renewable Energy Laboratory. 

http://www.nrel.gov/docs/fy07osti/41168.pdf. 

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