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

Gasoline from Wood via 

Integrated Gasification, 

Synthesis, and Methanol-toGasoline Technologies 

Steven D. Phillips, Joan K. Tarud, Mary J. Biddy, 

and Abhijit Dutta

Technical Report

NREL/TP-5100-47594

January 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

Gasoline from Wood via 

Integrated Gasification, 

Synthesis, and Methanol-toGasoline Technologies 

Steven D. Phillips, Joan K. Tarud, Mary J. Biddy, 

and Abhijit Dutta

Prepared under Task No. BB07.3710

Technical Report

NREL/TP-5100-47594

January 2011
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

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iii

Executive Summary

This report documents the National Renewable Energy Laboratory’s (NREL’s) assessment of the 

feasibility of making gasoline via the methanol-to-gasoline (MTG) route using syngas from a 

2,000 dry metric tonne/day (2,205 U.S. ton/day) biomass-fed facility. 

The thermochemical route of biomass gasification produces a syngas rich in hydrogen and 

carbon monoxide. The syngas is then converted into methanol, and the methanol is converted to 

gasoline using the methanol-to-gasoline (MTG) process first developed by Exxon Mobil. Using a 

methodology similar to that used in previous NREL design reports and a feedstock cost of 

$50.70/dry U.S. ton ($55.89/dry metric tonne), a plant gate price (PGP) was estimated. For the 

base case the PGP is predicted to be $16.60/MMBtu ($15.73/GJ) (U.S. $2007) for gasoline and 

liquefied petroleum gas (LPG) produced from biomass via gasification of wood, methanol 

synthesis, and the methanol-to-gasoline process (MTG). The corresponding unit prices for 

gasoline and LPG are $1.95/gallon ($0.52/liter) and $1.53/gallon ($0.40/liter) with yields of 55.1 

and 9.3 gallons per U.S. ton of dry biomass (229.9 and 38.8 liters per metric tonne of dry 

biomass), respectively. For comparison to ethanol, this is $1.39 per gallon ($0.37/liter) ethanol 

on an energy equivalent basis. In comparison, based on analysis work completed at NREL, the 

predicted plant gate prices for ethanol produced via the thermochemical and biochemical 

pathways are $1.57 per gallon ($0.41 per liter) and $1.49 per gallon ($0.39 per liter), respectively

(OBP 2009). Note that the PGP is for the base case. A sensitivity analysis is included in the 

report to demonstrate the impact that modifications in the design and costing assumptions have 

on the PGP. A range of PGP values is to be expected due to uncertainties in capital costs, yields, 

and technoeconomic factors. 

This report is a future look at the potential of the described biomass-to-gasoline process, based 

on calculations for a mature plant (also called the nth plant) and 2012 technology targets as 

established in the Multi-Year Technical Plan of the U.S. Department of Energy (DOE) Office of 

the Biomass Program. In order to achieve the $1.95/gallon ($0.52/liter) PGP, there are critical 

research milestones that must be achieved. First, the 2012 tar reforming targets of 99.9% tar and 

80% methane conversion (among others) are essential. Also, utilization of a fluidized bed MTG 

reactor, instead of the commercially proven fixed bed, is pertinent in keeping capital costs down. 

Thus, further research on this type of reactor is needed 1) to verify that at the conditions 

specified the products generated match the model assumptions and 2) to analyze effects of scaleup on product distribution. It should be emphasized that the PGP for a first-of-a-kind plant will 

be significantly higher than the PGP for an nth plant. 

To predict the PGP for this study, a new technoeconomic model was developed in Aspen Plus, 

based on the model developed for NREL’s thermochemical ethanol design report (Phillips et al. 

2007). The necessary process changes were incorporated into a biomass-to-gasoline model using 

a methanol synthesis operation followed by conversion, upgrading, and finishing to gasoline. 

Results of the simulation were used to obtain mass and energy flows, which were then used to 

size and estimate the cost of process equipment in an Excel spreadsheet-based economic model. 

This report follows the approach taken in the thermochemical ethanol design case: the DOE 

Office of the Biomass Program’s 2012 research targets were used for the gasifier and tar 

reformer operation (Phillips et al. 2007). The methanol and MTG processes were modeled using 

published results.
 

 

 

 

 

 

A discounted cash flow rate of return (DCFROR) calculation was performed to determine the 

PGP required to meet a 10% internal rate of return (IRR). A thermal basis approach was used to 

account for co-products (LPG and electricity). Instead of assigning a market value to co-products 

and then using income from the sale of those products to offset operating costs, we used total 

energy production to determine the PGP on a cost per energy basis (e.g., $/MMBtu or $/GJ) 

using all products in the calculation. The higher heating value of the individual products was 

then used to calculate the volumetric cost of the fuels and the per-kilowatt-hour cost of 

electricity. This approach allocates a proportional fraction of the capital and operating costs for 

the plant to each of the main products. 

The overall plant efficiency was 42.6% (lower heating value [LHV] basis) and the carbon 

efficiency to LPG and gasoline was 31%. The efficiency to the desired gasoline product was 

37.7% LHV and 28% carbon efficiency. The gasifier efficiency was 74.9%. Potential process 

improvements include utilizing more of the tail gases to make products other than heat and 

electricity. Because all of the power for the plant ultimately comes from the biomass fed to the 

plant, any energy efficiency improvements to the plant should improve product yields. 

iv 
v 

Table of Contents

Executive Summary................................................................................................................................... iii

Table of Contents........................................................................................................................................ v

List of Figures ............................................................................................................................................ vi

List of Tables .............................................................................................................................................. vi

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

1.1 Background ..........................................................................................................................2

1.2 Methanol and Methanol-to-Gasoline (MTG) Technology Discussion ................................4

1.3 Gasoline Discussion .............................................................................................................9

1.4 Analysis Approach .............................................................................................................10

1.5 Process Design Overview ..................................................................................................13

1.6 Feedstock and Plant Size ...................................................................................................15

2 Process Description........................................................................................................................... 16

2.1 Feed Handling & Preparation – Area 100 ..........................................................................17

2.2 Gasification – Area 200 .....................................................................................................18

2.3 Gas Cleanup & Conditioning – Area 300 ..........................................................................19

2.4 Methanol Synthesis – Area 400 .........................................................................................22

2.5 Methanol Conditioning – Area 500 ...................................................................................23

2.6 Methanol to Gasoline Conversion (MTG) – Area 1400 ....................................................24

2.7 Gasoline Separation – Area 1500 ......................................................................................25

2.8 Steam System and Power Generation – Area 600 .............................................................27

2.9 Cooling Water and Other Utilities – Area 700 ..................................................................29

2.10 Additional Design Information ........................................................................................31

2.11 Thermal and Pinch Analyses ............................................................................................31

2.12 Water Demands ................................................................................................................32

3 Process Economics ........................................................................................................................... 33

4 Economics - Results .......................................................................................................................... 38

4.1 Cost Contribution for Gasoline ..........................................................................................39

4.2 Sensitivity Analyses ...........................................................................................................40

5 Conclusions ........................................................................................................................................ 43

6 Acknowledgments.............................................................................................................................. 44

7 References .......................................................................................................................................... 45

Appendix A. List of Acronyms................................................................................................................. 50

Appendix B. Comparison of Aspen Model to Four MTG Compositions from Literature................... 51

Appendix C. NREL Biorefinery Design Database Description and Summary .................................... 54

Appendix D. Individual Equipment Cost Summary ............................................................................... 61

Appendix E. Economic Summary Page from Excel Spreadsheet........................................................ 70

Appendix F. Discounted Cash Flow Rate of Return Summary ............................................................ 71

Appendix G. Heat Exchanger Network ................................................................................................... 75

Appendix H. Process Flow Diagrams ..................................................................................................... 78

Appendix I. Comparison to other Biomass-to-Gasoline and Methanol-to-Gasoline 

Published Costs................................................................................................................................ 101

I.1 Comparison to “Techno-economic Analysis for the Conversion of Lignocellulosic 

Biomass to Gasoline via the Methanol-to-Gasoline (MTG) Process” by S. B. Jones and 

Y. Zhu from Pacific Northwest National Laboratory (PNNL) ........................................101

I.2 Comparison of Total Project Investment to Published Cost Information from the New 

Zealand MTG Commercial Plant (Seddon 2006) ............................................................102

Appendix J. External Reviewer Comments and Responses .............................................................. 103
vi 

List of Figures

Figure 1. Process schematic of the crude gasoline separations area ............................................... 8

Figure 2. Approach to process analysis ........................................................................................ 10

Figure 3. Chemical Engineering Magazine's Plant Cost Indices .................................................. 13

Figure 4. Block flow diagram ....................................................................................................... 13

Figure 5. Current case design block flow diagram of thermochemical gasoline from biomassderived methanol and the methanol-to-gasoline process ............................................... 17

Figure 6. Composition of conventional gasoline wt % (typical) .................................................. 27

Figure 7. Pinch analysis composite curves ................................................................................... 32

Figure 8. Cost breakdown by area in $/gallon .............................................................................. 40

Figure 9. Sensitivity analysis for biomass-to-gasoline process .................................................... 43

List of Tables 

Table 1. Gasifier and Tar Reformer Performance Targets in 2012 ................................................ 3

Table 2. Fluidized Bed Reactor Pressure (P1000) Effects on LPG Yields and PGP ...................... 5

Table 3. Fixed Bed Reactor Conditions for the Methanol to Gasoline Processes .......................... 6

Table 4. Gasoline Regulations from Before the Clean Air Act of 1990, as a Result of the Clean 

Air Act of 1990, and California Regulations for Reformulated Gasoline ......................... 9

Table 5. Chemical Engineering Magazine's Plant Cost Indices ................................................... 12

Table 6. Ultimate Analysis of Hybrid Poplar Feed ....................................................................... 16

Table 7. Gasifier Operating Parameters, Gas Compositions, and Efficiencies ............................ 19

Table 8. Current and Target Design Performance of Tar Reformer ............................................. 20

Table 9. Tar Reformer Conditions and Outlet Gas Composition ................................................. 20

Table 10. Acid Gas Removal Design Parameters ......................................................................... 21

Table 11. Process Conditions for Methanol Synthesis ................................................................. 23

Table 12. Composition of Crude Methanol Intermediate in Model .............................................. 24

Table 13. MTG Reaction Conditions and Yields .......................................................................... 25

Table 14. Composition of LPG from Aspen Model ..................................................................... 27

Table 15. Power Requirements for Plant by Process Area ........................................................... 29

Table 16. Utility and Miscellaneous Design Information ............................................................. 31

Table 17. Process Water Demands ............................................................................................... 33

Table 18. Installed Equipment Costs by Process Area ................................................................. 34

Table 19. General Cost Factors in Determining Total Installed Equipment Costs ....................... 34

Table 20. Cost Factors for Indirect Costs ..................................................................................... 34

Table 21. Breakdown of Operating Cost Contribution to PGP .................................................... 35

Table 22. Variable Operating Costs .............................................................................................. 36

Table 23. Economic Parameters ................................................................................................... 37

Table 24. Labor Costs ................................................................................................................... 38

Table 25. Other Fixed Costs ......................................................................................................... 38

Table 26. Process and Economic Results Summary for 2012 BTG Case .................................... 39

Table 27. List of Variables for Sensitivity Analyses .................................................................... 41

Table I-1. Comparison of NREL and PNNL Model Assumptions and Effect on PGP .............. 101
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

1 Introduction 

In his 2006 State of the Union Address, President Bush declared that America is “addicted to oil” 

and announced the Advanced Energy Initiative (AEI), which included increased research funding 

for cutting edge biofuels production processes. In response to the AEI, Congress passed a 

Renewable Fuel Standard (RFS) as part of the Energy Independence and Security Act (EISA) of 

2007 that requires 36 billion gallons (136.3 billion liters) of biofuels per year by 2022 (Biomass 

Research and Development Board 2008). In February 2010, President Obama and the U.S. 

Environmental Protection Agency (EPA) announced a finalized Renewable Fuel Standard 2.0 

(RFS2) to implement the long-term RFS (Biomass Intel 2010). In reference to the RFS, U.S. 

Department of Energy (DOE) Secretary Chu stated, “Developing the next generation of biofuels 

is key to our effort to end our dependence on foreign oil and address the climate crisis...” (The 

White House 2009). 

The current analysis was conducted to investigate one of several possible biofuels that can be 

produced using the thermochemical route of gasification and synthesis. The basis for this study 

was a stand-alone gasification/synthesis process including sub-processes or unit operations for 

integrated tar reforming, acid gas scrubbing, and synthesis to methanol followed by conversion 

to gasoline. This biomass-to-gasoline process will be referred to as the BTG process in this 

report. 

The starting point for this study was the model developed for the National Renewable Energy 

Laboratory’s (NREL’s) 2007 thermochemical ethanol design report (EDR) (Phillips et al. 2007). 

The report was based on achieving research targets in key barrier areas for cellulosic ethanol. 

That study concluded that, within the typical uncertainties inherent to this type of conceptual 

technoeconomic analysis, ethanol could be produced at a Minimum Ethanol Selling Price 

(MESP) of $1.01 per gallon ($0.27 per liter) in 2005 U.S. dollars based on a feedstock cost of 

$35 per dry U.S. ton ($38.60 per dry metric tonne). The DOE Office of the Biomass Program 

(OBP) Multi-Year Program Plan (MYPP) updated the MESP to $1.57 per gallon ($0.41 per liter) 

in 2007 U.S. dollars with a feedstock cost of $50.70 per dry U.S. ton ($55.89 per dry metric 

tonne) and with adjusted alcohol synthesis targets. As published in the EDR, the yield was 80.1 

gallons of fuel grade ethanol per U.S. ton of dry biomass (334.2 liters of ethanol per metric tonne 

of dry biomass). As published in the OBP MYPP, the ethanol yield is 71.1 gallons of ethanol per 

dry U.S. ton of biomass (296.7 liters of ethanol per dry metric tonne of biomass) because of 

adjusted alcohol synthesis targets (OBP 2009). 

The feedstock for the EDR was poplar wood with 50% moisture. The same feedstock is used 

here. Prior to being fed into the gasifier, the feedstock is dried to a moisture level between 5 wt 

% and 20 wt %, depending on the amount of waste heat available for drying. In the EDR, the 

wood moisture level was 5 wt % at the dryer outlet. In this study, the process conditions dictated 

a higher moisture level (10 wt %) because insufficient waste heat was available to achieve the 

same level of dryness as in the EDR. The gasifier efficiency was reduced due to this higher 

moisture content in the feed, which slightly affected the “raw” syngas composition. 

The front end of the thermochemical process is similar in both the production of ethanol (EDR) 

and the production of gasoline (via BTG). Because a new slate of products is being formed in the 

BTG case, thus using different catalytic processes, the overall process heat integration and 

1
 
 

 

 

 

 

 

 

 

 

 

 

 

materials requirements are modified. So, although the front-end process equipment is similar in 

both cases, the process flows and equipment sizes are significantly different between the two 

processes. In both studies, the best use of process and heat streams was attempted to achieve 

optimum productivity and economics. The complexity of the thermochemical processes makes it 

difficult to determine whether the best economic scenarios have been found. It is likely, 

especially in the BTG case, that better scenarios can be developed for achieving higher gasoline 

yields and lower production costs. 

1.1 Background 

Prior to the publication of the EDR in 2007, a 2005 milestone report (Aden et al. 2005) reviewed 

the history of thermochemical technoeconomic studies at NREL. Sections of that report, as well 

as some of the EDR, are repeated here for convenience and updated with work completed since 

those reports were published. 

An extensive literature search on mixed alcohols research and technology was included in the 

2005 milestone report (Aden et al. 2005). A technical evaluation firm was subcontracted to 

document the current state of mixed alcohols technology for NREL (Nexant 2006a-d). Several 

conceptual process designs and models were developed to generate detailed mass and energy 

balance data. NREL’s previous thermochemical design report (Spath et al. 2005) served as the 

basis for the feed preparation, drying and handling, gasification, gas cleanup and conditioning, 

and compression sections for the process model. For the EDR, ethanol was obtained by mixed 

alcohol synthesis, separation, and purification, and the appropriate sections were added to the 

model. The mass and energy balance data were used within a discounted cash flow rate of return 

economic analysis (DCFROR), along with capital and operating costs, to calculate the minimum 

product selling price required to meet a 10% internal rate of return (IRR). Sensitivity analyses 

were conducted around several parameters to determine the extent of their impact on the overall 

economics of the process. 

In 2006, the thermochemical models from previous studies were reviewed and updated. The 

detailed EDR (Phillips et al. 2007) was published in January 2007. One of the new elements of 

that report was that it looked only at the technoeconomic results for the year 2012 using the 

target performance goals established in the DOE OBP Multi-Year Technical Plan (OBP 2008). A 

second update was that the operating parameters that are the subject of OBP research targets 

were estimated for 2012. (The present BTG study uses the same values for these parameters, as 

shown in Table 1.) Third, in addition to the tar reformer targets used in the previous studies, 

which were kept unchanged, anticipated improvements in the mixed alcohol catalyst were 

incorporated into the EDR study. The sensitivity to the target parameters was shown along with 

other non-research parameters used in the study such as feedstock moisture and cost. 

2
 
 

 

 

 

 

 

 

 

 

 

 

 

Table 1. Gasifier and Tar Reformer Performance Targets in 2012 

Compound 2012 Targets for Tar and Methane 

Conversion Reforming 

Methane (CH4) 80% 

Ethane (C2H6) 99% 

Ethylene (C2H4) 90% 

Tars (C10+) 99.9% 

Benzene (C6H6) 99% 

Ammonia (NH3)

a 90% 

a Converts to N2 and H2 

A lack of published data on mixed alcohol catalysts – and non-existent published data on 

commercial mixed alcohol catalysts – significantly increased the uncertainty of the EDR results, 

because some potential performance parameters had to be deduced from other similar catalyst 

systems with published results (e.g., Fischer-Tropsch and methanol catalysts). The key result of 

the EDR was that fuel ethanol could be produced from biomass via gasification and mixed 

alcohols synthesis for $1.01 per gallon ($0.27 per liter) (based on a feedstock cost of $35 per 

U.S. ton [$38.60 per metric tonne] and in 2005 dollars), a price slightly below the cost target 

defined in 2005 ($1.07 per gallon ethanol [$0.28 per liter ethanol]). 

Based on the EDR, it became possible to address questions on different technologies and fuels 

using a common base model and common assumptions, where meaningful. The first derivative 

of the EDR was a report which looked at a different gasifier design (Dutta and Phillips 2008). 

The EDR used an indirectly heated steam gasifier operating at near-atmospheric pressure. Dutta 

and Phillips incorporated many of the features of the EDR model into a new model using an 

oxygen-blown, medium-pressure gasifier to produce syngas that was later converted into ethanol. 

New equipment, such as an air separation unit (ASU), was added to the design and economic 

evaluation. That study showed that the alternate gasifier design would not achieve the $1.07 per 

gallon ($0.28 per liter) minimum cost target for ethanol. The additional cost of the ASU and 

lower methane conversions in the reformer under higher operating pressures were major 

contributors to the higher MESP of $1.57 per gallon ($0.41 per liter) (using the 2005 cost 

assumptions, thus comparing with $1.01 per gallon [$0.27 per liter] ethanol via the indirect 

gasification process). 

Besides ethanol, investigated in the EDR, many other fuels have been investigated over the past 

30 years. Those include methanol, Fischer-Tropsch liquids (FTL), and gasoline from the 

methanol-to-gasoline (MTG) process (Bartholomew and Farrauto 2006). All of these 

investigations used processes producing fuels on a large scale using syngas derived from 

abundant sources, such as coal and natural gas. In the latter case, excess or stranded natural gas 

was converted into methanol, allowing the gas to be made transportable to end users. As a result, 

in all areas of the world except for China, methanol is produced almost entirely from natural gas 

reforming (steam methane reforming) (Haddeland 1981). FTL has been used successfully in 

South Africa for more than 50 years to provide gasoline to that country. There, syngas produced 

3
 
 

 

 

 

 

 

 

 

 

 

 

 

via coal gasification is fed to catalytic reactors, which were initially developed by several 

companies; however, Sasol’s proprietary technology in this field became dominant later. 

In the late 1970s, Mobil Oil (now part of ExxonMobil) developed the MTG process to convert 

natural gas via methanol into gasoline (Schreiner 1978). The gasoline was called M-gasoline. A 

technically successful demonstration plant was operated in New Zealand for several years before 

being dismantled because of the inability to compete against lower priced petroleum-derived 

gasoline. More recently, in 2007 DKRW Advanced Fuels began work to create a coal-togasoline plant using ExxonMobil’s MTG process. The plant is located in Carbon County, 

Wyoming, and is predicted to come online in 2013 with an initial production capacity of 15,000– 

20,000 barrels (630,000–840,000 gallons or 2.4 million–3.2 million liters) per day (DKRW 

Advanced Fuels 2007). Also in 2008, Synthesis Energy Systems announced its agreement with 

ExxonMobil for up to 15 methanol-to-gasoline licensed plants in their global operations, the first 

of which would produce 7,000 barrels (294,000 gallons or 1.1 million liters) per day and be 

located near Benwood, West Virginia (BusinessWire 2008). Due to the economic downturn in 

2008, the latter plant could not be financed. Though not modeled here, another method, TIGAS, 

has been developed from the MTG technology to produce gasoline from synthesis gas in a 

single-loop process, thus eliminating the need for methanol production and storage (Haldor 

Topsoe 2010). 

A natural extension of the aforementioned ethanol fuel studies is to combine the biomass-tosyngas technologies from the EDR with the other existing syngas conversion technologies 

developed over nearly 80 years of research. The purpose of this report is to provide one such 

analysis: the BTG process. 

Perhaps the main advantage of the biomass-to-gasoline process is that it produces a “drop-in 

fuel,” i.e., a fuel that can be accommodated in the current motor-fuel infrastructure without any 

adaptations to it. It should also have lower risks for investors than processes with less developed 

technologies, because the methanol synthesis process is well developed and commercial. The 

MTG process has commercial experience, with proven performance using conditioned syngas, 

especially when using the fixed bed MTG reactors. The risk in the front end of the process, 

upstream from the synthesis reactor(s), is the same for both the thermochemical ethanol and 

gasoline processes because they are essentially identical. 

At the plant gate, the point at which this evaluation ends, the potential benefit to a gasoline 

product is that it can be shipped long distances using the interstate pipeline system already in 

place with few, if any, restrictions. The fuel is fungible with the existing fuel distribution 

infrastructure, although it may need to be blended with more conventional gasoline products to 

meet current fuel specifications. 

1.2 Methanol and Methanol-to-Gasoline (MTG) Technology Discussion 

A plethora of published literature exists for syngas conversion to methanol, with bibliographies 

available in books by Lee (Lee 1990) and Bartholomew (Bartholomew and Farrauto 2006). The 

first-generation methanol synthesis catalyst, developed by BASF in 1923, required a temperature 

of 300°C–400°C and pressure of up to 30 MPa. Since that time, the required pressure (4–100 

MPa) and temperature (180°C–250°C) have decreased. A commercial methanol synthesis 

catalyst today can have selectivity to methanol as high as 99.9% and yields of up to 2.28 kg of 

4
 
 

 

 

 

 

 

 

 

 

 

 

methanol per liter of catalyst per hour. Because the methanol catalysts are commercially 

available, the synthesis reaction should typically perform as presented in the literature. No 

discussion of methanol reaction mechanisms or other highly technical information is given here. 

Interested readers are referred to the references used above. 

In contrast to the abundant literature for methanol synthesis, there is a dearth of published 

literature for the methanol-to-gasoline (MTG) process. The MTG process was developed by 

Mobil (now part of ExxonMobil) in the late 1970s. A DOE report from 1978 assessing gasoline 

production from coal using MTG and Sasol-type Fischer-Tropsch technologies provided much of 

the information used in the present study (Schreiner 1978). The DOE study was done in part by 

researchers from Mobil Research and Development Corporation. Several references (Probstein 

and Hicks 1982; Edwards and Avidan 1986; Chang 1992; Gary and Handwerk 1994; Mokrani 

and Scurrell 2009) also provided limited insight to the operating conditions, product 

distributions, and process equipment needs. 

The 1978 DOE report is one of the earliest published studies on the MTG process. It was based 

on a Lurgi high-pressure coal gasifier to generate syngas and the Lurgi methanol synthesis 

process to make methanol for the MTG process. The MTG data used in the report came from a 

jointly funded process development study, using fixed bed reactors, between Mobil and DOE 

under contract E (49-18)-1773 (Schreiner 1978). 

In the MTG process the methanol, or rather its dehydrated derivative dimethylether (DME), is 

reacted over a ZSM-5 zeolite catalyst, on which the chain growth of molecules is sterically 

hindered, thus allowing only production of gasoline and lighter material. This report evaluates 

the MTG process for a fluidized bed reactor, in which case direct conversion is possible, because 

both the conversion of methanol to DME as well as the synthesis of DME to gasoline can be 

performed in one reactor. The fluidized bed reactor has been technically proven at the 

demonstration plant in Germany (100–200 bpd [4,200–8,400 gal/day; 15,900–31,800 L/day]). 

The running conditions for a typical fluidized bed MTG reactor are 400°C and 60 psi (413 kPa) 

(Probstein and Hicks 1982; Gayubo et al. 1999). However, in this study the MTG reactor 

pressure is 200 psi (1.38 MPa). The higher pressure improved LPG recovery and yields in the 

model. Table 2 shows the LPG yields and PGP as a function of MTG reactor pressure. These 

changes in LPG yields do not include possible product composition changes with respect to 

pressure. 

Table 2. Fluidized Bed Reactor Pressure (P1000) Effects on LPG Yields and PGP 

P1000 Pressure 

(psig) (MPa) 

PGP 

($/gallon) ($/liter) 

LPG Yield 

(lb/h) (kg/h) 

200 1.38 

150 1.04 

100 0.69 

90 0.62 

60 0.41 

$1.95 $0.515 

$1.96 $0.518 

$1.98 $0.523 

$2.00 $0.528 

$2.17 $0.573 

3,659 1,663 

3,426 1,557 

2,982 1,355 

2,791 1,269 

1,354 615 

The fluid bed scale-up to pilot scale from bench scale (100–200 bpd [4,200-8,400 gal/day; 

15,900-31,800 L/day] from 4 bpd [168 gal/day; 636 L/day]) was reported to have been achieved 

5 
 

 

 

 

 

 

 

 

 

 

 

without loss of conversion efficiency while the heat management, steady activity level, higher 

gasoline yield, and octane number were retained. The successful scale-up is attributed to 

attention to Peclet number, a superficial velocity greater than 0.3 m/s, and turbulent flow through 

the reactor. The demonstration plant in Germany was reported to have conversion efficiencies 

greater than 99.9% (Edwards and Avidan 1986). 

However, if a fixed bed reactor is desired, the conversion process will need to take place in two 

steps. The first step is to pass the methanol over a methanol dehydration catalyst to form a 

mixture of methanol, DME, and water, and the second step is converting this mixture to gasoline 

over the ZSM-5 zeolite catalyst. In the 1977 report, a large gas recycle was used to limit the 

temperature rise across the ZSM-5 fixed catalyst bed to 125°F (52°C). The ZSM-5 zeolite 

catalyst is known to deactivate via coke formation over a period of about two weeks. In the fixed 

bed case, five parallel reactors are used to convert the DME/methanol mixture, with one reactor 

being taken offline every two weeks for regeneration with air to remove the coke. In the 

fluidized bed reactor, the catalyst is continuously withdrawn and regenerated by partially burning 

off the coke (Mokrani and Scurrell 2009). 

Typical reactor conditions for each of the two fixed bed conversion steps are given in Table 3 

(Schreiner 1978). 

Table 3. Fixed Bed Reactor Conditions for the Methanol to Gasoline Processes 

Reactor Conditions (Schreiner) 

First Reactor (Methanol to DME) 

Pressure, inlet 401 psig (2.77 MPa)
 

Temperature, inlet 680°F (360°C)
 

Temperature, outlet 788°F (420°C)
 

Space velocity 6 lb fresh feed/h/lb catalyst
 

Estimated catalyst life 2 years
 

Second Reactor (DME to Gasoline) 

Pressure, inlet 200 psig (1.38 MPa)
 

Temperature, inlet 625°F (330°C)
 

Temperature, outlet 752°F (400°C)
 

Space velocity 1.84 lb fresh feed/h/lb catalyst
 

Estimated catalyst life 1 year
 

The MTG process dehydrates methanol to DME and subsequently synthesizes DME to gasoline 

hydrocarbons and LPG with stoichiometric yields to hydrocarbons and water of 44% and 56%, 

respectively (Schreiner 1978). According to the 1978 DOE report, “The gasoline is chemically 

conventional consisting of highly branched paraffins (51%), highly branched olefins (13%), 

naphthenes (8%), and aromatics (28%).” Amongst the latter, one less desirable product, 1,2,4,5­

tetramethylbenzene (durene) is formed. An aromatics content of 28% is high, but the gasoline is 

envisioned to be added to a very large gasoline stream with low aromatic content. The gasoline 

yields modeled in this report are based upon the results published in the 1978 DOE report. 

However, because the large recycle of light hydrocarbons necessary for the fixed bed case is not 

necessary in the fluidized bed scenario, the gasoline product spectrum will contain more light 

6
 
 

 

 

 

 

 

 

 

 

 

 

 

 

olefins and fewer light paraffins (Mokrani and Scurrell 2009). MTG gasoline is also reported to 

have a Research Octane Number (RON) between 90 and 100, with no products created with a 

carbon number greater than 10 and no oxygenates (Schreiner 1978). 

Gasoline fuel specifications have changed significantly since 1978, in part because of the Clean 

Air Act of 1990 (as amended). Of particular issue are 1) the low benzene concentration (< 1.0 

vol %) allowed in today’s fuel and 2) the lower Reid Vapor Pressure (RVP) required in most 

parts of the country (7.0–7.8 psi [48–54 kPa] vs. 10 psi [69 kPa] in 1978). Lower RVP 

requirements are met today, in part, by removing most of the butanes and butenes from the 

gasoline. The gasoline composition will be discussed further in Section 2.2. 

The hydrocarbon product from the MTG process has more than 51 compounds, similar to 

straight-run gasoline in a petroleum refinery. The 51 compounds reportedly made in the MTG 

process (Schreiner 1978) are listed in Appendix B along with product compositions from two 

other literature sources. The several sources agree on the composition for most of the 

compounds. The compositions are given in varying degrees of detail with different methods used 

to lump groups of compounds. 

The process for upgrading of the gasoline mixture is similar to the process used in a gasoline 

refinery. The design used in this study came from the New Zealand MTG demonstration process 

design with a few minor modifications. Figure 1 shows the process schematic for the crude 

gasoline separations area. 

The first separation step is to remove ethane from the “crude” gasoline stream using a distillation 

column known as a de-ethanizer. The “de-ethanized” gasoline from the bottom of the column 

goes to a second column known as a stabilizer that removes the lighter components, propanes 

and propenes (3-carbon hydrocarbons referred to as C3s), and most of the butanes and butenes. 

These light components are also known in the industry as liquefied petroleum gas (LPG), which 

can be sold or further processed to increase the gasoline yield as discussed below. 

The amount of butanes and other light hydrocarbons is limited in gasoline because of RVP 

restrictions. To increase the gasoline yield, while utilizing the butanes and other lighter 

hydrocarbons, refineries typically use an alkylation unit. Alkylation is a process that joins a 

hydrocarbon containing a tertiary carbon (a carbon attached to three other carbons) with an olefin 

to create a larger branched molecule from the two smaller molecules. Of particular interest is the 

reaction of isobutane with 2-butene to form isooctane. If instead of 2-butene, 1-butene is reacted 

with isobutane, 2-methylheptane is formed. The MTG gasoline compositions given in Schreiner 

1978 and in Liederman 1978 do not distinguish between 1- and 2-butene but lump them as 

“butenes.” Thus specific yields of 1-butene and 2-butene are unknown. The Aspen simulation is 

set for the alkylation unit to produce isooctane from isobutane and 2-butene. If the product is 

primarily 1-butene, possible reconfiguration to include an isomerization unit would shift the 

product yield to include greater 2-butene levels. 

The addition of the hydrofluoric acid (HF) alkylation unit improves the utilization of C4 

hydrocarbons while making a higher octane product, isooctane. Other alkylation products are 

possible using pentene and an isoparaffin, but typically these are not upgraded because they are 

suitable to the gasoline product as is and would require additional distillation. C3 hydrocarbons 

7
 
 

 

 

can also be upgraded to larger molecules in the gasoline range. Because the alkylation step 

requires an isoparaffin to react with an olefin, additional steps, like isomerization, can be used to 

convert n-butane (straight-chained) to the iso form, which can then be reacted with olefins. The 

most common commercial alkylation processes are catalyzed by a strong acid, either 

hydrofluoric acid or sulfuric acid. While environmental and safety concerns with respect to the 

use of strong acid catalysts have prompted development of solid acid catalyst methods for 

alkylation units, only traditional alkylation units are considered in this study due to performance 

and cost data availability (Hutson 1977). 

The unreacted C3 and C4 hydrocarbons from the alkylation unit are either separated from the 

heavier gasoline products in another column and sent to a storage tank for sale as LPG or 

returned to the stabilizer column where the alkylate joins the gasoline stream. If a separate 

column is used to separate the C3 and C4 hydrocarbons, then the alkylate is sent to a tank for 

storage and later blended with the other gasoline fractions. 

The bottoms product from the stabilizer is sent to a splitter column to separate the stabilized 

gasoline into light and heavy fractions. The light fraction exits the top of the column where it is 

condensed and sent to a storage tank. The heavy fraction could be sent to an isomerization 

reactor to convert 1,2,4,5-tetramethylbenzene (durene) into 1,2,3,5-tetramethylbenzene. The 

former product (durene) has a relatively high freezing point (77°F, 25°C) and will crystallize in 

fuel systems if the concentration exceeds about 5 vol %. The isomer product has a lower freezing 

point. This study concluded that the isomerization reactor was unnecessary because of the 

dilution effect of adding the BTG gasoline to a very large conventional gasoline stream. 

Figure 1. Process schematic of the crude gasoline separations area 

8
 
 

 

 

 

 

 

 

 

 

 

1.3 Gasoline Discussion 

In 1990, the Clean Air Act mandated the seasonal use of oxygenated compounds in gasoline in 

specific regions of the United States. According to a Government Accountability Office (GAO) 

report (June 2005), in 2005 there were approximately 45 different gasoline blends in use across 

the United States. The large number of products was due, in part, to the approaches taken to 

improve air quality throughout the country. Each state can voluntarily pass laws specifying the 

gasoline blends that must be used to meet air quality standards, and many states have done so. 

The EPA has required some states to use reformulated gasoline blends to meet air quality 

standards in cases where those standards have not been met. Each state can propose the method 

and fuels it wants to use to meet the air quality standards. The EPA must approve any gasoline 

formulation proposed by the states that meets the criteria for achieving the improved air quality 

specifications. 

Table 4. Gasoline Regulations from Before the Clean Air Act of 1990, as a Result of the Clean Air 

Act of 1990, and California Regulations for Reformulated Gasoline 

Pre-1990 1990 Clean Air Act California Air 

Resources Board 

Benzene 2% 1% max 1.0 vol % max 

Oxygen 0.20% 2% min 1.8–2.0 mass % 

Sulfur 150 ppm 150 ppm max 40 ppm 

Aromatics 32.00% 25% max 25 vol % max 

Olefins 9.90% 5% max 6 vol % max 

RVP 8.7 psi (60 kPa) 7.3/8.1 psi (50/56 kPa) (south/north) 7.0 psi (48 kPa) 

90% evap. 170°C NA 149°C 

Table 4 lists key characteristics for reformulated gasolines. Two characteristics are typically 

specified for improving the ozone-producing potential of gasoline: the RVP and the oxygenate 

content. Decreased vapor pressure specifications required refiners to remove the more volatile 

compounds, mainly butanes and butenes, from gasoline. The maximum amount of n-butane is 

added to the final gasoline product while still meeting the RVP limit, which varies by season and 

local temperatures. The use of oxygenates was implemented to improve the gasoline combustion 

in car engines and to decrease the amount of unburned hydrocarbons, carbon monoxide, and 

nitrous oxides emitted from the tailpipe. The oxygenate requirement is typically met by “splash 

blending” ethanol into gasoline being loaded into tanker trucks for delivery to filling stations. 

Gasoline that will have oxygenate added at the terminal must be blended at the refinery gate to 

account for any changes in RVP and octane ratings. 

With the various combinations of oxygenates and RVP requirements, multiplied by the consumer 

desire for three octane products to be available at each filling station, the number of products that 

must be managed in the gasoline distribution infrastructure has proliferated. It is possible that the 

gasoline product from an MTG process will be sent to an existing petroleum refinery via 

pipeline, blended into the refinery’s standard gasoline products, and then shipped again via 

pipeline or truck to the final blending stations where additives and oxygenates (ethanol) are 

splash blended before final distribution to filling stations. It is also possible that the MTG fuel 

could be sent to distributors and blended with conventional gasoline to meet final product 

9
 
 

 

 

 

 

 

 

 

 

 

 

 

specifications before being sent on to filling stations, but that would require an additional storage 

tank for the MTG product and a new methodology for blending gasoline at a location other than 

the refinery. 

1.4 Analysis Approach 

The general approach used in the development of the process design, process model, and 

economic analysis is depicted in Figure 2 and described below (previously described in Phillips 

et al. 2007). The first step was to assemble a general process flow schematic and more detailed 

process flow diagrams (PFDs). (See Appendix H for the associated PFDs for this design.) From 

this, detailed mass and energy balance calculations were performed around the process using 

Aspen Plus software. Data from the Aspen model were then used to properly size all process 

equipment and fully develop an estimate of capital and operating costs. These costs could 

potentially be used in several types of economic analyses. For this design, a discounted cash flow 

rate of return (DCFROR) analysis was used to determine the Plant Gate Price (PGP) necessary to 

meet a small economic profitability or internal rate of return (IRR) of 10%. 

Process Flow DOE/NREL Diagrams Sponsored 

Research Results 

Rigorous Material 

& Energy Balances 

(Aspen Plus) 

Outside 

Engineering 

Studies 

Vendor Cost 

Quotations 

Capital & Project 

Cost Estimation 

Estimates of 

Other 

Commercial 

Technology 

Cost Estimation 

Software 

(e.g. Aspen IPE) 

Engineering 

Company Cost 

Estimates Discounted Cash 

Flow Model 

Minimum Gasoline 

Selling Price 

Figure 2. Approach to process analysis 

This thermochemical conversion process was developed based upon in-house experience 

performing conceptual designs for biomass conversion to ethanol via biochemical means (Aden 

et al. 2002), biopower applications, and biomass gasification for hydrogen production (Spath et 

al. 2005). Specific information for potential sub-processes was obtained through a subcontract 

with Nexant Inc. (Nexant 2006a-d). 

10
 
11 

Aspen Plus version 2004.1 was used to determine the mass and energy balances for the process. 

The operations were separated into nine major HIERARCHY areas (the numbering gap between 

Areas 700 and 1400 is there to keep consistency with past reports; Areas 800 through 1300 were 

entirely different sections not applicable to this process): 

• Feed Handling and Drying (Area 100) 

• Gasification (Area 200) 

• Cleanup and Conditioning (Area 300) 

• Methanol Synthesis (Area 400) 

• Methanol Conditioning (Area 500) 

• Steam Cycle (Area 600) 

• Cooling Water (Area 700) 

• Methanol-to-Gasoline (Area 1400) 

• Gasoline Separation (Area 1500) 

Overall, the Aspen Plus simulation consists of about 400 operation blocks (such as reactors and 

flash separators), 1,070 streams (720 mass-, 265 heat-, and 85 work-streams), and 80 control 

blocks (design specifications and calculator blocks). Many of the gaseous and liquid components 

were described as distinct molecular species using Aspen’s component properties database. The 

raw biomass feedstock, ash, and char components were modeled as non-conventional 

components. There was more detail and rigor in some blocks (e.g., distillation columns) than 

others (e.g., conversion extent in the methanol synthesis reactor). Because this design processes 

three phases of matter (solids, gases, and liquids), no single thermodynamics package was 

sufficient. Instead, multiple thermodynamics packages were used in the Aspen simulation as 

needed to model the various process streams and unit operations. The RKS-BM option was used 

throughout much of the process for high temperature, high pressure phase behavior. The Aspen 

default steam tables, STEAM-TA, were used for the steam cycle calculations, and the 

ELECNRTL package was used to model the electrolyte species potentially present within the 

quench water system.

The process economics are based on the assumption that this is the “nth” plant, meaning that 

several plants using this same technology will have already been built and are operating. This 

means that additional costs for risk financing, longer start-ups, and other costs associated with 

first-of-a-kind plants are not included. 

The capital costs were developed from a variety of sources. For some sub-processes that are well 

known technologies and can be purchased as modular packages (i.e., amine treatment, acid gas 

removal), an overall cost for the package unit was used. Many of the common equipment items 

(tanks, pumps, simple heat exchangers) were costed using the Aspen Icarus Process Evaluator 

and Aspen Questimate costing software. For other more specific unit operations (gasifier, 

LOCAT system), cost estimates from other studies and/or from vendor quotes were used. As 

documented in the hydrogen design report (Spath et al. 2005), the installed capital costs were 

developed using general plant-wide factors. The installation costs incorporated cost contributions 

not only for the actual installation of the purchased equipment but also for instrumentation and 

controls, piping, electrical systems, buildings, yard improvements, etc. These are also described 

in more detail in Section 3, and additional information is available in Appendices C and D. 
12 

The purchased component equipment costs reflect the base case for equipment size and cost 

year. The sizes needed in the process may actually be different than what was initially designed. 

Instead of re-costing in detail, an exponential scaling expression was used to adjust the base

equipment costs: 

( ) New Size New Cost Base Cost

Base Size

  =    

n

where n is a characteristic scaling exponent (typically in the range of 0.6 to 0.7). The sizing 

parameters are based upon some characteristic of the equipment related to production capacity, 

such as inlet flow or heat duty in a heat exchanger (appropriate if the log-mean temperature 

difference is fairly similar). Generally these related characteristics are easier to calculate and give 

nearly the same result as resizing the equipment for each scenario. The scaling exponent n can be 

inferred from vendor quotes (if multiple quotes are given for different sizes), multiple estimates 

from IPE or Questimate at different sizes, or a standard reference (such as Garrett 1989; Peters 

and Timmerhaus 2003; or Perry et al. 1997). It is known that for very large equipment, n can rise 

to almost 1.0 as economies of scale almost disappear and multiplication of equipment is 

inevitable. For the size of the equipment used in this evaluation, however, economy of scale is 

realizeable.

Because a variety of sources was used, the base equipment costs were derived based upon 

different cost years. Therefore, all capital costs were adjusted with the Chemical Engineering

(CE) magazine’s Plant Cost Index to a common basis year of 2007: 

( ) Cost Index in New Year New Cost Base Cost

Cost Index in Base Year

  =    . 

The CE indices used in this study are listed in Table 5 and depicted in Figure 3. The indices were 

very nearly the same for 2000 to 2002 (essentially zero inflation) but take a fairly sharp increase 

after 2003 (primarily due to increased worldwide steel prices). 

Table 5. Chemical Engineering Magazine's Plant Cost Indices

Year Index

2000 394.1

2001 394.3

2002 395.6

2003 402.0

2004 444.2

2005 468.2

2006 499.6

2007 525.4
 

 

 

 

 

 

 

550
 

500
 

450 

400 

350 

1999 2000 2001 2002 2003 2004 2005 2006 2007 2008 

Figure 3. Chemical Engineering Magazine's Plant Cost Indices 

Once the scaled, installed equipment costs were determined, we applied overhead and 

contingency factors to determine a total plant investment cost. That cost, along with the plant 

operating expenses (generally developed from the Aspen Plus model’s mass and energy balance 

results), was used in a discounted cash flow analysis to determine the cost of gasoline production 

(referred to as the plant gate price or PGP). For the analysis done here, the PGP is the primary 

value used to compare alternate designs. 

1.5 Process Design Overview 

Figure 4. Block flow diagram 

A simple block flow diagram of the current design is depicted in Figure 4. The detailed process 

flow diagrams (PFDs) are in Appendix H. The process has the following steps (the process steps 

up to the methanol synthesis section were previously explained in Phillips et al. [2007], with 

modifications made to the gas cleanup and conditioning section): 

•	 Feed Handling & Preparation. The biomass feedstock (2,000 dry metric tonne/day 

[2,205 dry U.S. ton/day]) is dried from the as-received moisture content to that required 

for proper feeding into the gasifier using flue gases from the char combustor and tar 

13
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

reformer catalyst regenerator. Prior to drying, wood chips with a diameter larger than 2 

inches are sent to the hammer-mill for further size reduction. 

•	 Gasification. Indirect gasification is considered in this assessment. Heat for the 

endothermic gasification reactions is supplied by circulating hot synthetic olivine1 “sand” 

between the gasifier and the char combustor. Conveyors and hoppers are used to feed the 

biomass to the low-pressure indirectly-heated entrained flow gasifier. Steam is injected 

into the gasifier to aid in stabilizing the entrained flow of biomass and sand through the 

gasifier. The biomass is chemically converted to a mixture of syngas components (CO, 

H2, CO2, CH4, etc.), tars, and a solid “char” that is mainly the fixed carbon residual from 

the biomass plus carbon (coke) deposited on the sand. Cyclones at the exit of the gasifier 

separate the char and sand from the syngas. These solids flow by gravity from the 

cyclones into the char combustor. Air is introduced to the bottom of the combustor 

reactor and serves as a carrier gas for the fluidized bed plus as the oxidant for burning the 

char and coke. The heat of combustion heats the sand to more than 1800°F (982°C). The 

hot sand and residual ash from the char is carried out of the combustor by the combustion 

gases and separated from the hot gases using another pair of cyclones. The first cyclone is 

designed to capture mostly sand while the smaller ash particles remain entrained in the 

gas exiting the cyclone. The second cyclone is designed to capture the ash and any sand 

passing through the first cyclone. The hot sand captured by the first cyclone flows by 

gravity back into the gasifier to provide the heat for the gasification reaction. Ash and 

sand particles captured in the second cyclone are cooled, moistened to minimize dust, and 

sent to a landfill for disposal. 

•	 Gas Cleanup & Conditioning. This consists of multiple operations: reforming of tars and 

other hydrocarbons to CO and H2; syngas cooling/quench; and acid gas (CO2 and H2S) 

removal. Tar reforming is envisioned to occur in an isothermal fluidized bed reactor; deactivated reforming catalyst is separated from the effluent syngas and regenerated online. 

The hot syngas is cooled through heat exchange with a steam cycle and additional 

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

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

wastewater treatment facility. In order to increase methanol production, a sulfur resistant 

low-temperature-water-gas-shift (LTS) reactor improves the syngas H2:CO ratio after 

scrubbing. After leaving the LTS, the syngas enters an amine unit for removal of the CO2 

and H2S and subsequently enters the methanol synthesis reactor. The H2S is reduced to 

elemental sulfur and stockpiled for disposal. The CO2 is vented to the atmosphere in this 

design. 

•	 Methanol Synthesis. The cleaned and conditioned syngas is converted to methanol in a 

fixed bed reactor containing a copper/zinc oxide/alumina catalyst. The mixture of 

methanol and unconverted syngas is cooled through heat exchange with the steam cycle 

and other process streams. The methanol is separated by condensing it away from the 

unconverted syngas. Unconverted syngas is recycled back to the entrance of the methanol 

synthesis reactor. 

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

used as a sand for various applications. A small amount of magnesium oxide (MgO) is added to the fresh olivine to 

prevent the formation of glass-like bed agglomerations that would result from biomass potassium interacting with 

the silicate compounds. 

14
 
 

 

 

 

 

 

 

 

 

 

 

 

•	 Methanol Conditioning. The methanol leaving the reactor has been condensed at elevated 

pressure and has absorbed a sizeable quantity of gas. The methanol and gas stream is first 

heated and sent through a turbo expander generator to recover a portion of the 

compression energy. Once the stream is at a lower temperature it is sent to a distillation 

column to degas the methanol. This removal of gases could be done at a later stage in the 

process. 

•	 Methanol-to-Gasoline. The methanol is then passed through a fluidized bed reactor 

containing the ZSM-5 zeolite catalyst. Direct conversion to gasoline is achieved in the 

fluidized bed reactor. The gasoline product from the MTG process has more than 51 

compounds, similar to straight-run gasoline in a petroleum refinery. 

•	 Gasoline Separation. The separation of the gasoline mixture is similar to the process used 

in a gasoline refinery. The design used in this model came from the New Zealand MTG 

demonstration process design with a few minor modifications, as shown in Figure 1. This 

design utilizes five distillation columns to separate the remaining gas, LPG, light 

gasoline, and heavy gasoline. The remaining gas is sent to the fuel combustor. The light 

gasoline continues without further treatment. And the heavy gasoline could proceed 

through a durene isomerizer in order to eliminate the presence of the 1,2,4,5­

tetramethylbenzenes by converting them to 1,2,3,5-tetramethylbenzenes. This stream 

would then be merged with the light gasoline. The two product streams are LPG and 

gasoline. 

•	 Heat & Power. A conventional steam cycle produces heat (as steam) for the gasifier and 

reformer operations and electricity for internal power requirements (with the possibility 

to export excess electricity as a co-product). The steam cycle is integrated with the 

biomass conversion and MTG processes. Pre-heaters, steam generators, and super-heaters 

are integrated within the process design to create the steam. The steam will run through 

turbines to drive compressors, generate electricity, or be withdrawn at various pressure 

levels for injection into the process. The condensate will be sent back to the steam cycle, 

de-gassed, and combined with makeup water. 

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

of each cooling water heat exchanger within the biomass conversion process as well as the 

requirements of the cooling tower. 

Previous analyses of gasification processes have shown the importance of properly utilizing the 

heat from the high temperature streams. A pinch analysis was performed to analyze the energy 

network of this gasoline production process. Details of the pinch analysis will be discussed in 

Section 3.11. 

1.6 Feedstock and Plant Size 

Based upon expected availability per the “Billion Ton” vision study (Perlack et al. 2005), forest 

resources were chosen as the primary feedstock. The “Billion Ton” study addressed short- and 

long-term availability issues for biomass feedstocks without giving specific time frames. In the 

target year of 2012, it is most likely that only the “existing” and “unexploited” resources can be 

counted on to supply a thermochemical processing facility. Therefore, it is logical to base 

thermochemical processing on the forest resources. Thermochemical processing could provide a 

15
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

cost-effective technology to process this major portion of the expected biomass feedstock 

(Phillips et al. 2007). 

The design plant size for this study–—2,000 dry metric tonne/day (2,205 dry U.S. ton/day)—was 

chosen to match that of the biochemical process (Aden et al. 2002). For the process described 

here, the plant would produce just over one million barrels per year or 42.5 million gallons 

(160.9 million liters) per year. With an expected 8,406 operating hours per year (96% operating 

factor or stream factor), the annual feedstock requirement is 700,000 dry metric tonne/yr 

(772,000 dry U.S. ton/yr). 

The delivered feedstock cost was chosen to match recent analyses done at Idaho National 

Laboratory to target $50.70 per dry U.S. ton ($55.89 per dry metric tonne) by 2012 (OBP 2009). 

Cost effects due to feedstock cost were also examined as part of the sensitivity analysis. 

Past analyses have used hybrid poplar wood chips delivered at 50 wt % moisture to model forest 

resources; the same will be done here. The ultimate analysis for the feed used in this study is 

given in Table 6. Performance and cost effects due to composition and moisture content were 

examined as part of the sensitivity analysis and alternate scenarios (Phillips et al. 2007). 

Table 6. Ultimate Analysis of Hybrid Poplar Feed 

Component (wt %, dry basis )a 

Carbon 50.88 

Hydrogen 6.04 

Nitrogen 0.17 

Sulfur 0.09 

Oxygen 41.90 

Ash 0.92 

Heating valueb (Btu/lb) 8,671 HHV (20.1 MJ/kg) 

8,060 LHV (18.7 MJ/kg) a Craig and Mann 1996. b Calculated using the Aspen Plus Boie correlation. 

2 Process Description 

As mentioned above, the starting point for this model was the thermochemical ethanol model 

used in the EDR. The synthesis reactor was changed to make predominantly methanol, and the 

syngas conditioning requirements (e.g., sulfur and CO2 concentration, H2:CO ratio) and reaction 

conditions (temperature, pressure, residence time) were changed to match the methanol process 

requirements. The post-synthesis sections of the process were a major addition to the EDR 

model. Although the fluidized bed MTG process design was given in other reports, there were 

enough differences in the current model to prevent direct use of those process parameters in this 

study. Specifically, since no outside energy was allowed in this design, the lightest hydrocarbons 

were combusted instead of proceeding to the alkylation unit. However, butanes still proceed to 

the alkylation unit. Distillation column operating parameters were set by using column design 

specifications that would give “acceptable” effluent compositions. Because gasoline does not 

have a specific composition but rather a range of acceptable compositions for various component 

groups (e.g., aromatics, olefins, paraffins) that meet overall performance criteria and physical 

16
 
       

   

   

characteristics such as octane number and vapor pressure, an attempt to select reasonable design 

specifications was used. 

A block-flow diagram depicting the current case design is shown in Figure 5 for a 2,000 dry 

metric tonne/day (2,205 dry U.S. ton/day) BTG process. The front end of the process (through 

steam reforming) remains substantially similar to the EDR (Phillips et al. 2007). 

Process flow diagrams (PFDs) for the BTG process are available in Appendix H. 

Qu

Feed Prep

Dryer

Gasification

and

Cyclones

Fluidized Bed

Tar Reformer

Quench and

Scrubber

FlueGas

rted

as

t

t

ed

Raw

Syngas

Wet,

“Tar Free”

Syngas

Dry

Syngas

d

&

Char

Combustion

and Cyclones

Gas

Combustion

& Catalyst

Regeneration

Syngas

Compression

Water Gas

Biomass

Flue Gas

Diver

Syng

Regenerat

Catalyst

UsedCatalyst

HotSand

CharSand

Ash

Regeneration Water‐Gas

Shift

CO2/H2S

Removal

Methanol

Synthesis

Syngas/

Methanol Degassing

Electricity

Steam

Cycle

Unreacted/Recycled Syngas

Methanol Synthesis Clean Removal

Separation

MTG

Conversion

Gasoline

Finishing

Dissolved

Gases

Clean

Syngas

Gasoline

LPG

CarbonDioxide

SolidSulfur

Methanol,

Unreacted

Syngas

Methanol

Crude

Gasoline

LPG

Figure 5. Current case design block flow diagram of thermochemical gasoline from biomassderived methanol and the methanol-to-gasoline process 

2.1 Feed Handling & Preparation – Area 100 

This section of the process accommodates the delivery of the biomass feedstock, short-term onsite storage, and the preparation of the feedstock for the gasifier. The design is based upon a 

woody feedstock. It is expected that the feed handling area for agricultural residues would be 

very similar. 

The feed handling and drying sections are shown in PFD-P850-A101 and PFD-P850-A102. 

Wood chips are delivered to the plant primarily via trucks; delivery by train could be an 

attractive alternative. As the trucks enter the plant they are weighed (M-101), and the wood chips 

are dumped into a storage pile. From the storage pile, the wood chips are conveyed (C-102) 

through a magnetic separator (S-101) and screened (S-102). Particles larger than 2 inches are 

sent through a hammer-mill (T-102/M-102) for further size reduction. Front end loaders transfer 

the wood chips to the dryer feed bins (T-103). 

17 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Drying is accomplished by direct contact of the biomass feed with hot flue gas. The 2,000 dry 

metric tonne/day (2,205 dry U.S. ton/day) plant requires two identical, parallel feed handling and 

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

screw conveyor (C-104). The wood is dried to a moisture content of 10 wt % with flue gas from 

the char combustor (R-202) and the tar reformer’s fuel combustor (R-301). The exhaust gas 

exiting the dryer is sent through a cyclone (S-103) and baghouse filter (S-104) to remove 

particulates prior to being emitted to the atmosphere. The stack temperature is controlled by 

cooling the hot flue gas from the char combustor and the tar reformer with two steam boilers (H286B and H-311B) prior to entering the dryer. This generated steam is added to the common 

steam drum (T-604) (see Section 3.8). The dried biomass is then conveyed to the gasifier train 

(T-104/C-105). 

2.2 Gasification – Area 200 

This section of the process converts a mixture of dry feedstock and steam to syngas and char 

(also described in Phillips et al. 2007). Heat is provided in an indirect form by circulating olivine 

that is heated by the combustion of the char downstream of the gasifier. The steam primarily acts 

as a fluidizing medium in the gasifier and also participates in certain reactions when high gasifier 

temperatures are reached. 

From the feed handling and drying section, the dried wood enters the gasifier section as shown in 

PFD-P850-A201. The 2,000 dry metric tonne/day (2,205 dry U.S. ton/day) plant was modeled 

using two parallel gasifier trains. The gasifier (R-201) used in this analysis is a low-pressure 

indirectly-heated circulating fluidized bed (CFB) gasifier. The gasifier was modeled using 

correlations based on run data from the Battelle Columbus Laboratory (BCL) 9 metric tonne/day 

(9.9 U.S. ton/day) test facility (Bain 1992). 

The heat for the endothermic 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, a calcined magnesium silicate, primarily Enstatite (MgSiO3), Forsterite (Mg2SiO3), and 

Hematite (Fe2O3). A small amount of magnesium oxide (MgO) must be added to the fresh 

olivine because it titrates the potassium in the feed ash. Without the MgO addition, the potassium 

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

(approximately 930°F, 500°C), and its formation will cause the bed media to become sticky, 

agglomerate, and eventually defluidize. Adding MgO makes the potassium form a high melting 

point (approximately 2,370°F, 1,300°C) ternary eutectic with the silica, thus sequestering it. 

Potassium carryover in the gasifier/combustor cyclones is also significantly reduced. The ash 

content of the feed is assumed to contain 0.2 wt % potassium. The MgO flow rate is set at 2 

times the molar flow rate of potassium. 

The gasifier fluidization medium is steam supplied from the steam cycle (see Section 3.8). The 

steam-to-feed ratio is 0.4 lb of steam per lb of dry biomass. The gasifier pressure is 23 psia (159 

kPa). The olivine circulating flow rate is 27 lb of olivine per lb of dry wood. Fresh olivine is 

added at a rate of 0.01% of the circulating rate to account for losses. The char combustor is 

operated with 20% excess air. 

Both the gasifier and char combustor temperatures are dictated from the energy balances around 

the gasifier and combustor. The gasifier temperature is 1,622°F (883°C) and the char combustor 

18
 
 

 

 

 

 

 

 

 

temperature is 1816°F (991°C). The composition of the outlet gas from the gasifier is shown in 

Table 7. 

Table 7. Gasifier Operating Parameters, Gas Compositions, and Efficiencies 

Gasifier Variable Value 

Temperature 1,622°F (883°C) 

Pressure 23 psia (159 kPa) 

Gasifier Outlet Gas Composition mol % (wet) mol % (dry) 

H2 13.9 24.7 

CO2 7.1 12.6 

CO 23.7 42.0 

H2O 43.6 -­

CH4 8.6 15.2 

C2H2 0.2 0.4 

C2H4 2.4 4.2 

C2H6 0.1 0.2 

C6H6 0.07 0.1 

Tar (C10H8) 0.1 0.2 

NH3 0.2 0.3 

H2S 0.04 0.1 

H2:CO molar ratio 0.59 

Stoichiometric ratio	 1.047 

Gasifier Efficiency	 75.3% HHVa basis
 

74.9% LHVb basis
 a Higher Heating Value. b Lower Heating Value. 

2.3 Gas Cleanup & Conditioning – Area 300 

This section of the process cleans and conditions the syngas so that the gas can be synthesized 

into methanol. In Area 300, the tars and hydrocarbons in the syngas are reformed to additional 

CO and H2. Particulates are removed by quenching. Acid gases (CO2 and H2S) are removed, and 

the syngas is compressed. 

The gas from the secondary gasifier cyclone is sent to the catalytic tar reformer (R-303), shown 

in PFD-P850-A301. In this fluidized bed reactor the hydrocarbons are converted to CO and H2 

while NH3 is converted to N2 and H2. In the Aspen simulation, the conversion of each compound 

is set to match targets that are believed to be attainable through near-term research efforts. Table 

8 gives the experimental conversions (for deactivated catalyst) that have been achieved at NREL 

(Phillips et al. 2004; Dutta and Aden 2008) and the conversions used in the simulation 

corresponding to the 2012 research targets. Section 4.2 includes PGP information if the tar 

reformer conversions from Phillips et al. 2004 are used. The composition of the gas leaving the 

tar reformer in the Aspen simulation is shown in Table 9. 

19
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 8. Current and Target Design Performance of Tar Reformer 

Compound Experimental Conversion 

to CO and H2 

Target Conversion 

to CO and H2 

Methane (CH4) 50% 80% 

Tars (C10+) 99.6% 99.9% 

Benzene (C6H6) 97.9% 99% 

Table 9. Tar Reformer Conditions and Outlet Gas Composition 

Tar Reformer Variable Value 

Tar reformer inlet temperature 

Tar reformer outlet temperature 

1,622°F (883°C) 

1,622°F (883°C) 

Tar Reformer Outlet Gas Composition mol % (wet) mol % (dry) 

H2 45.26 54.55 

CO2 8.05 9.70 

CO 27.48 33.12 

H2O 17.03 -­

CH4 1.42 1.71 

C2H2 0.02 0.024 

C2H4 0.18 0.22 

C2H6 9.74 ppmv 11.74 ppmv 

C6H6 4.56 ppmv 5.50 ppmv 

Tar (C10H8) 0.83 ppmv 1.00 ppmv 

NH3 0.01 0.012 

H2S 0.03 0.036 

N2 0.52 0.63 

H2:CO molar ratio 1.65 

Stoichiometric number - 1.04 

(H2-CO2)/(CO+CO2) 

The hot syngas is cooled through heat exchange with the steam cycle (H-301A-H) and with 

cooling water via scrubbing, shown in PFD-P850-A302. The scrubbing system consists of a 

venturi scrubber (M-302) and a quench chamber (M-301). It removes impurities such as 

particulates and ammonia along with any residual tars. The scrubbing system quench water is a 

closed recirculation loop with heat rejected to the cooling tower and a blowdown rate of 

approximately 82.4 gpm (311.9 L/min) sent to a wastewater treatment facility. Any solids that 

settle out in T-301 are sent off-site for treatment as well. 

The steam reformer has a significant water-gas-shift potential because of its nickel-based 

catalyst. On a single pass system at NREL, the H2:CO ratio has reached 4:1 under some 

operating conditions. This extent of WGS is not necessarily the best scenario for making 

methanol because it also produces 1 mole of CO2 for every mole of H2 made. The CO2 in the 

syngas must be removed to achieve the specified level of 5 vol % at the synthesis reactor. Carbon 

utilization to the desired final product can be improved by recycling unreacted syngas back to the 

synthesis reactor inlet. However, inerts in the gas limit the amount of gas that can be recycled, 

especially with the CO2 limitation. Purging a portion of the recycle stream is used to reduce the 

buildup of inerts, but it also slightly decreases the available syngas utilization. 

20
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

The quench step cools the syngas to a temperature of 140°F (60°C). The syngas is then 

compressed using a five-stage centrifugal compressor (K-301) with interstage cooling as shown 

in PFD-P850-A303. The compressor was modeled such that each section has a polytropic 

efficiency of 78% and an intercooler outlet temperature of 140°F (60°C). The interstage coolers 

are forced air heat exchangers. The syngas leaving the compressor is at 750 psi (5.2 MPa) 

(Phillips et al. 2007). 

A low temperature shift (LTS) process (R-434) was added to the model after syngas 

compression, as shown in PFD-P850-A401. A design specification in the Aspen Plus model was 

used to divert only as much syngas as needed to meet a H2:CO ratio of 2.1 (mole basis). The 

stream to the LTS was mixed with enough superheated steam at 900°F (482°C) to give a 

steam:CO ratio of 1.0 (mole basis). The LTS effluent has an H2:CO ratio of 6.38 and a CO2 mole 

percentage of 23.3%. The low temperature shift is completed in a single stage. It contains a 

copper based catalyst that contains reactive zinc oxide, which traps sulfur in the top of the bed as 

zinc sulfide and prevents sulfur poisoning (Twigg 1996). 

To meet the various conditions and expectations for product yields, a process with various 

recycle streams was designed. An amine-based CO2 removal step (acid gas removal or AGR) 

was left in the process design from the EDR model, with similar operating conditions and energy 

requirements, as shown in Table 10. The AGR separator (S-310) is shown in PFD-P850-A304. 

The gas is sent to AGR just before entering the methanol synthesis reactor to ensure that the gas 

entering the synthesis reactor is at the accepted levels of CO2 and H2S. 

Table 10. Acid Gas Removal Design Parameters 

Acid Gas Removal Parameter Value 

Amine used Monoethanolamine (MEA) 

Amine concentration 35 wt % 

Amine circ. rate 2,261.5 gpm (8,559.8 L/min) 

Amine temp. @ absorber 110°F (43.3°C) 

Absorber pressure 735 psia (5.1 MPa) 

Stripper condenser temperature 212°F (100°C) 

Stripper reboiler temperature 230°F (110°C) 

Stripper pressure 65 psia (449 kPa) 

Stripper reboiler duty 162 MMBtu/h (171 GJ/h) 

Stripper condenser duty 108 MMBtu/h (114 GJ/h) 

Amine cooler duty 54.3 MMBtu/h (57.3 GJ/h) 

Heat duty per pound CO2 removed 2,650 Btu/lb (6.19 MJ/kg) 

CO2 removed 61,170 lb/h (27,746 kg/h) 

CO2 removed in the scrubber is vented to the atmosphere. Prior to CO2 removal, the syngas 

stream needs to be compressed. The higher pressure improves the amine-CO2 equilibrium and 

gives better performance and lower energy requirements. A pressure of 735 psia (5.1 MPa) is 

used. The syngas must be quenched to remove any condensable material, primarily steam, prior 

to the compression step. 

21 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

The acid gases removed in the amine scrubber are then stripped to regenerate the sorbent and 

sent through a sulfur removal operation using a liquid phase oxidation process shown in PFDP850-A305. The combined amine/LO-CAT process will remove the sulfur and CO2 to the levels 

desired for the copper/zinc oxide/alumina catalyst. Although there are several liquid-phase 

oxidation processes for H2S removal and conversion available today, the LO-CAT process was 

selected because of its progress in minimizing catalyst degradation and for its environmentallybenign catalyst. LO-CAT is an iron chelate-based process that consists of a venturi precontactor 

(M-303), liquid-filled absorber (M-304), air-blown oxidizer (R-301), air blower (K-302), 

solution circulation pump (P-303), and solution cooler (H-305). The air flow rate for re-oxidizing 

the LO-CAT solution was included in the simulation and calculated based on the requirement of 

2 moles O2 per mole H2S. Prior to entering the LO-CAT system, the gas stream is superheated in 

the preheater (H-304) to 10°F (5.6°C ) above its dew point, which in this process is equivalent to 

148°F (64.4°C). This degree of superheating is required for the LO-CAT system. The LO-CAT 

process was modeled to remove the H2S to a concentration of 10 ppmv, which is the permissible 

exposure limit (University of Wisconsin 2007) in the CO2 vent effluent from the amine scrubber. 

The CO2 from the LO-CAT unit is vented to the atmosphere (Phillips et al. 2007). 

The specified limit for sulfur in the syngas was set at 0.1 ppmv, as per the literature. We assumed 

that this level could be achieved with the AGR removal system because it is also suitable to H2S 

removal. In practice, a ZnO guard bed would likely be used to protect the synthesis catalyst. 

Omitting the ZnO guard bed is not expected to impact the production cost calculated in this 

study. A sensitivity analysis that included the cost of the ZnO guard bed and the ZnO catalyst 

was run and concluded $0.00 increase to the PGP (set-up for ZnO guard bed found in Shumake 

and Small 2006). 

2.4 Methanol Synthesis – Area 400 

The cleaned and conditioned syngas is converted to methanol in a fixed bed reactor (R-490) 

containing a copper/zinc oxide/alumina catalyst, shown in PFD-P850-A403. The mixture of 

methanol and unconverted syngas is cooled through heat exchange with the steam cycle and 

other process streams. The liquid methanol is recovered by condensing it (H-411-414) and 

separating the liquids from the residual syngas (S-414). Almost 87 wt % of the unconverted 

syngas is recycled back to the entrance of the synthesis reactor (Hamelinck and Faaij 2002). This 

is one difference between the BTG process and the EDR method, which had no direct recycle to 

the synthesis reactor. 

To maintain consistency with the EDR, the same type of synthesis reactor was used in this study. 

However, multiple reactors in series with interstage cooling could be an alternative reactor 

configuration that may increase methanol yields. Table 11 lists the methanol synthesis conditions 

for a typical copper/zinc oxide/alumina catalyst given in Bartholomew (Bartholomew and 

Farrauto 2006). The methanol reaction is fast, and equilibrium is quickly achieved. The heat 

released during the reaction is a serious concern, because even short excursions of only a few 

degrees can seriously damage the catalyst irreparably. For this level of study, a kinetics model 

was not warranted; the temperature was assumed to be isothermal. A REQUIL model was used 

in the Aspen Plus model. The literature notes that CO2 concentrations of up to 7 vol % can 

improve productivity to methanol (Lee 1990). Excessive levels of CO2 decrease conversion. A 

value of 5 vol % CO2 was chosen because that was the level used in the EDR. 

22
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 11. Process Conditions for Methanol Synthesis 

“State of Technology” Conditions Used in Process 

Parameter Conditions Design & Aspen Model 

Temperature (°C) ~ 300 (572°F)a 300 (572°F) 

Pressure (psia) 735 (5.1 MPa)

a 735 (5.1 MPa) 

H2:CO ratio 2b 2.1 

CO2 concentration (mol %) 3%–8%a 5.0% 

Sulfur concentration (ppmv) < 0.1a 0.09 

Stoichiometric Number ­ 2b 1.73 

(H2 -CO2)/(CO+CO2) 

H2/(2CO+3CO2) 1.05a 0.87 a Bartholomew & Farrauto 2006. b Olah et al. 2006. 

The operating pressure to make methanol is significantly lower than that for making mixed 

alcohols (735 psia [5.1 MPa] vs. 1,000–2,000 psia [6.9–13.8 MPa]) (Phillips et al. 2007). The 

temperature is comparable to other synthesis reactions. The desire to have a stoichiometric 

number of approximately 2 with a concomitant CO2 concentration of 5%, along with a desire to 

maximize fuel production (vs. making electricity), poses a design challenge to find the best 

economic conditions. 

Boiler feed water was assumed to be cross-exchanged within the reactor to generate steam for 

the process. Other reactor designs (e.g., slurry bubble reactors) were not considered at this time, 

but they do merit investigation because they are reported to have good heat management 

characteristics. 

While productivity was not specified within the Aspen model, methanol synthesis catalysts have 

been reported to have productivity values of more than 1,000 g/L-cat/h, with 99% or better 

selectivity to methanol. Given the high selectivity reported in the literature, no other byproducts 

were assumed in this design, especially because any byproducts are reported to be converted by 

the downstream MTG process, though there could be exceptions. 

2.5 Methanol Conditioning – Area 500 

The vapor-phase product from the synthesis reactor must be cooled to recover the methanol and 

to allow unconverted syngas and any inert gaseous species (CO2, CH4) to be recycled or purged. 

Cooling water is used to lower the temperature to 90°F (32°C) (H411-414), a temperature at 

which a majority of the liquid methanol condenses and is separated in a knock-out vessel (S471). About 2,600 lb/h (1,180 kg/h) of methanol, or 3% of the total methanol, is not recovered 

from the product stream at this temperature. 

The methanol is still at elevated pressure at this point in the process, resulting in a significant 

quantity of gas being absorbed in the methanol stream as it leaves the synthesis section of the 

process. These gases are removed from the methanol at this stage of the process and are then 

recycled to the reformer inlet. It may be possible to remove the gases at a later stage in the 

process. Removal at this point in the process was for modeling convenience, because the stream 

could be mixed back into the unreacted syngas stream from the synthesis condensation train. The 

combined gas streams are heated (H-505) before expansion through a turbo expander generator 

(K-501) to recover some of the compression energy of the gas by generating electricity, shown in 

23
 
 

 

 

 

 

 

 

 

 

 

 

 

PFD-P850-A502. About 4% of this recycled syngas and other dissolved gases are purged into the 

fuel gas stream to prevent an accumulation of inert gases. The degassed methanol product is sent 

to a storage tank (T-592) for short-term surge buffering between the synthesis and MTG sections 

of the plant. 

The composition of this methanol product is shown in Table 12. After the degassing step, the 

product is nearly 96% methanol, with the remainder being mainly CO2 and water and small 

amounts of various components. 

Table 12. Composition of Crude Methanol Intermediate in Model 

NREL Model (vol %) 

Methanol 95.9% 

CO2 2.3% 

H2O 1.4% 

Others 0.4% 

2.6 Methanol to Gasoline Conversion (MTG) – Area 1400 

In the methanol-to-gasoline (MTG) process, methanol is reacted over a ZSM-5 zeolite catalyst. 

Prior to conversion, the crude methanol from the intermediate storage tank is pumped into the 

MTG process to raise the liquid pressure to 200 psia (1.4 MPa). The methanol is then passed 

over the ZSM-5 zeolite catalyst in a fluidized bed reactor (R-1410), as shown in PFD-P850­

1401. The reactor has a riser, a disengaging vessel, and cyclones located above the fluidized bed. 

In the fluidized bed reactor, the catalyst is continuously withdrawn and regenerated by partially 

burning off some of the coke (Mokrani and Scurrell 2009).The regenerator is a combustor type in 

which carbon deposited on the catalyst is burned off by an upward air stream. The catalyst 

returns via a slide valve. Therefore, no additional reactors are necessary, in contrast with the 

fixed bed case. No provisions for catalyst regeneration were considered in this stage of the MTG 

process. Table 13 lists the reaction conditions and yields for the MTG process. 

24
 
 

 

 

 

 

Table 13. MTG Reaction Conditions and Yields 

Temperature °C (°F) 

Reactor inlet 330 (625) 

Reactor outlet 400 (752) 

Pressure MPa (psia) 

Reactor inlet 1.45 (210) 

Reactor outlet 1.28 (185) 

Yield wt % 

Hydrocarbons 44 

Water 56 

Total 100 

Crude Hydrocarbon Product wt % 

Light gas 2 

Propane 5 

Propylene 1 

Isobutane 7 

n-Butane 5 

Butenes 1 

C5+ gasoline 79 

Total 100 

Finished Fuel Products wt % 

Gasoline 82 

LPG 10 

Fuel gas 8 

Total 100 

The Aspen model did not include all 51 MTG process products listed in the literature (see 

Appendix B). Key components such as benzene, durene, and light hydrocarbons were included 

because their specific fate is important to the final product quality and process heat integration. 

In the gasoline boiling range above C5 hydrocarbons, some isomers were lumped together 

because they will all be included in the final product and will stay together through the 

fractionation steps. Butanes and butenes were included because they are on the edge of 

acceptability for the gasoline product; too much of the C4 hydrocarbons can cause the RVP to 

exceed allowable limits. 

2.7 Gasoline Separation – Area 1500 

The separation of the gasoline mixture is similar to a typical gasoline refinery finishing section, 

as shown in Figure 1 and in the following PFD’s: PFD-P850-1501, -1502, and -1503. The design 

used in this model came from the New Zealand MTG demonstration process design with a few 

modifications. Other designs are possible. 

The first separation in this section is to remove the lighter hydrocarbons from the gasoline (deethanizer/de-propanizer) stream. This is done in the de-ethanizer column (D-1503). The 

25
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

overhead from this column contains about 50 wt % C4+ hydrocarbons, which need to be 

recovered. The overhead goes to an absorber column (D-1502) that uses lean oil from a 

downstream column as the absorbing liquid. The lean oil and absorbed hydrocarbons are fed 

back to the de-ethanizer column (D-1503) as reflux to strip out any light hydrocarbons captured 

in the absorber bottoms effluent. 

The bottoms product from the de-ethanizer (D-1503) is sent to a stabilizer column (D-1504) to 

remove the butanes. The column is operated to remove most of the butanes from the gasoline, 

thus “stabilizing” it. The overhead product is sent to an alkylation unit (R-1505), and the bottoms 

product is sent to a splitter column (D-1505) to split the stabilized gasoline into light and heavy 

gasoline fractions. A side draw from the splitter is sent to the absorber (D-1502) to provide 

reflux liquid for that column. 

The bottoms of the gasoline splitter, consisting mainly of higher hydrocarbons and aromatics, 

could be sent to an isomerization reactor (R-1500) to convert 1,2,4,5-tetramethylbenzene 

(durene) into 1,2,3,5-tetramethylbenzene. The former product (durene) has a relatively high 

freezing point and will crystallize in fuel systems if the concentration exceeds about 5 vol %. 

The isomer product has a lower freezing point. The effluent from the isomerization reactor 

would be sent to tank storage. The 1978 DOE report states that the isomerizer uses a small 

amount of hydrogen (Schreiner 1978). An MTG report from Pacific Northwest National 

Laboratory (PNNL) accounted for this hydrogen by adding a pressure swing adsorption unit after 

the methanol synthesis reactor (Jones and Zhu 2009). Other reports have concluded that the 

isomerizer is not necessary because the MTG gasoline will be added to a very large gasoline 

stream, thus decreasing the durene concentration to an acceptable level. This is the approach 

taken in this report. If the isomerizer is added to the process, there is no change in cost to the 

PGP. 

The isomeric mixture of butanes and butenes from the stabilizer is sent to an HF alkylation unit 

(R-1505) to convert isobutane and butene into isooctane. The effluent from the alkylation unit is 

sent to a final LPG/alkylate splitter column (D-1506) to separate the unreacted C4 hydrocarbons 

from the isooctane. Refrigeration is used to cool the LPG/alkylate splitter. The bottoms product 

is sent to storage. The final composition of the MTG gasoline from the BTG process is compared 

to a typical conventional gasoline and to the reported composition for “M-gasoline” from the 

DOE report in Figure 6 (compositions are also shown in Appendix B). The overhead mixture of 

butanes is considered a separate product – LPG. The final composition of the LPG is shown in 

Table 14. 

26
 
 

 

 

 

 

100% 

90% 

80% 

70% 

60% 

50% 

40% 

30% 

20% 

10% 

0% 

25.0 

0.0 

23.4 

52.2 

11.2 

6.6 

28.6 

10.0 

15.6 

10.0 

51.8 55.0 

Conventional Gasoline Aspen Model Gasoline M-Gasoline 

Paraffins Olefins Cycloparaffins Aromatics 

Figure 6. Composition of conventional gasoline wt % (typical)
 

Table 14. Composition of LPG from Aspen Model
 

Compound Name mol % 

Propane 28.6 

Propene 4.2 

i-Butane 42.8 

n-Butane 24.4 

Water 11 ppmv 

The LPG stream could be used to make more gasoline product by using existing technologies to 

isomerize the n-butane into isobutane and reacting the isobutane with lighter olefins now being 

sent to the fuel combustor as fuel gas. Alternatively, LPG is a marketable product and could be 

sold as is, or propane, also a marketable product, could be sold if it is recovered from the fuel 

gas. 

2.8 Steam System and Power Generation – Area 600 

This process design includes a steam cycle that produces steam by recovering heat from the hot 

process streams throughout the plant. Steam demands for the process include the gasifier, amine 

system reboiler, LO-CAT preheater, and gasoline separation distillation columns. Of these, only 

27
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

the steam to the gasifier and the steam to the low temperature shift are directly injected into the 

process; the rest of the plant heat demands are provided by indirect heat exchange of steam with 

process streams that have condensate return loops. Power for internal plant loads is produced 

from the steam cycle using an extraction steam turbine/generator (M-602). Power is also 

produced from the process expander (K-412), which takes the unconverted syngas from 563 psia 

(3.88 MPa) to 34 psia (234 kPa) before it is recycled to the tar reformer. Steam is supplied to the 

gasifier from the low pressure turbine exhaust stage. The plant energy balance is managed to 

generate only the amount of electricity required by the plant. The steam system and power 

generation area is shown in PFD-P850-A601, -A602, and -A603 in Appendix H. 

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

the process reboilers along with the steam turbine condensate and makeup water. The total 

condensate stream is heated to the saturation temperature and sent to the deaerator (T-603) to 

remove any dissolved gases out of the water. The water from the deaerator is first pumped to a 

pressure of 497 psia (3.43 MPa) (P-604) and then pre-heated to its saturation (bubble point) 

temperature using a series of exchangers. The saturated steam is collected in the steam drum (T604). To prevent solids buildup, water must be periodically discharged from the steam drum. The 

blowdown rate is equal to 2% of the water circulation rate. The saturated steam from the steam 

drum is superheated with another series of exchangers. The superheated steam temperature and 

pressure were set as a result of thermal analysis. Superheated steam enters the turbine sequence 

(M-602A, B, C) at 900°F (482°C) and 472 psia (3.25 MPa) and is expanded to a pressure of 176 

psia (1.21 MPa). The remaining steam then enters the low pressure turbine and is expanded to a 

pressure of 65 psia (448 kPa). Here a slipstream of steam is removed and sent to the gasifier and 

other exchangers. Finally, the steam enters a condensing turbine and is expanded to a pressure of 

1 psia (6.89 kPa). The steam is condensed in the steam turbine condenser (H-601), and the 

condensate is returned to the condensate collection tank. 

This model assumes that all drives for pumps, fans, etc. are electric motors. Table 15 contains the 

power requirement of the plant broken out into the different plant sections. Syngas compression 

is the largest power requirement for the plant (totaling 22,300 kW, or approximately two-thirds 

of the plant’s power demands). The plant power demands and power production were 

specifically designed to be nearly equal. Therefore, no excess power is being sold to or 

purchased from the grid. This plant was designed to be as energy self-sufficient as possible. This 

was accomplished by burning a portion of the “dirty” unreformed syngas in the fuel combustor 

(Section 300). While this does have a negative impact on the overall yields of the process, it 

negates the purchase of natural gas or grid power. 

28
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 15. Power Requirements for Plant by Process Area 

Plant Area Work, kW 

Feed handling & drying 742 

Gasification 3,447 

Tar reforming, cleanup & conditioning 25,645 

LTS, methanol synthesis and degassing 1,993 

MTG conversion 57 

Gasoline finishing 126 

Steam system and power generation 306 required 

34,346 generated 

Cooling water and other utilities 1,044 

Total plant power requirement 33,360 

2.9 Cooling Water and Other Utilities – Area 700 

The cooling water system is shown in PFD-P850-A701. A mechanical draft cooling tower (M701) provides cooling water to several heat exchangers in the plant. The tower utilizes large fans 

to force air through circulated water. Heat is transferred from the water to the surrounding air by 

the transfer of sensible and latent heat. Cooling water is used in the following pieces of 

equipment with the associated water demands: 

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

gasifier/combustor (698 lb/h [317 kg/h])
 

•	 The quench water recirculation cooler (M-301), which cools the water used in the syngas 

quench step (6,622 lb/h [3,004 kg/h]) 

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

the syngas after the last stage of compression (19,829 lb/h [8,994 kg/h]) 

•	 The LO-CAT absorbent solution cooler (H-305), which cools the regenerated solution 

that circulates between the oxidizer and absorber vessels (28 lb/h [12.7 kg/h]) 

•	 The reacted syngas cooler (H-414), which cools the gas in order to condense out the 

liquid methanol (10,025 lb/h [4,547 kg/h]) 

•	 The post methanol expander cooler, which cools the methanol prior to degassing (H504B) (11,846 lb/h [5,373 kg/h]) 

•	 The post methanol degassing cooler (H-593) (375 lb/h [170 kg/h]) 

•	 Post MTG reactor cooling (H-1414) (803 lb/h [364 kg/h]) 

•	 The cooling for various operations conducting gasoline separation (H-1500) (3,178 lb/h 

[1,442 kg/h]), including: 

o	 Cooling needed by the absorber column (D-1502) 

o	 Refrigeration cooling needed by the alkylate/LPG splitter column (D-1506) 

o	 Finishing cooler for LPG (H-1591) 

29
 
 

 

 

 

 

 

 

 

 

 

 

o	 Finishing cooler for gasoline (H-1593) 

•	 The cooler for the side draw of the gasoline splitter, which provides reflux liquid to the 

absorber column (H-1512B) (167 lb/h [76 kg/h]) 

•	 The blowdown water-cooled cooler (H-603), which cools the blowdown from the steam 

drum (2,961 lb/h [1,343 kg/h]) 

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

turbine (118,810 lb/h [53,891 kg/h). 

Makeup water for the cooling tower is supplied at 14.7 psia (101 kPa) and 60°F (16°C). Water 

losses 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 water supply. Evaporation losses 

and blowdown were calculated based on information and equations in Perry et al. (1997). The 

cooling water is supplied to the process at a pressure of 65 psia (448 kPa) and temperature of 

80°F (27°C) (Liptak 2005). It returns to the cooling tower at a temperature of 110°F (43°C). 

Refrigeration is utilized where cooling is needed below 90°F (32°C). A centrifugal compressor 

refrigeration system was selected. An instrument air system is included to provide compressed 

air for both service and instruments. The instrument air system is shown in PFD-P850-A701. 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 (792 kPa) 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: 

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

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

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

an MgO lock hopper (T-706) 

•	 An ammonia storage tank (T-704) and pump (P-704) 

•	 A hydrazine storage tank (T-707) and pump (P-705) for oxygen scavenging in the
 

cooling water.
 

This equipment is shown in PFD-P850-A702. 

30
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

2.10 Additional Design Information 

Table 16 contains some additional information used in the Aspen Plus model and the production 

design. 

Table 16. Utility and Miscellaneous Design Information 

Item Design Information 

Ambient air conditionsa,b,c Pressure: 14.7 psia (101 kPa) 

TDry Bulb: 90°F (32°C) 

TWet Bulb: 80°F (27°C) 

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 (13.8 kPa) 

Heat exchangers and packed beds = 5 psi (34.5 kPa) a In the GPSA Engineering Data Book (GPSA 1987), see Table 11.4 for typical design values for dry bulb 

and wet bulb temperature by geography. Selected values would cover summertime conditions for most of
 

the lower 48 states.
 b In Weast (1981), see F-172 for composition of dry air. Nitrogen value was adjusted slightly to force mole
 

fraction closure using only N2, O2, Ar, and CO2 as air components.
 c In Perry et al. (1997), see psychrometric chart, Figure 12-2, for moisture content of air.
 

2.11 Thermal and Pinch Analyses 

Thermal and pinch analyses were performed to analyze the energy exchanges throughout the 

plant. Energy integration is tremendously important to the overall efficiency and economics of 

the process. Therefore, a detailed understanding of how and where the energy is utilized and 

recovered is a necessity. The pinch technique was used as a systematic method for confirming 

that no thermodynamic laws were in violation with the modeled energy integration for the 

processes. While some heat integration was included, a thorough heat integration optimization 

was not completed; there is still great room for improvement. 

In order to do the pinch analysis, temperature and heat duty data were gathered for streams 

needing heating or cooling throughout the process. This information was input into Linhoff 

March SuperTarget software and was utilized to develop the composite curves (temperature vs. 

enthalpy graph) shown in Figure 7 and to create the Heat Exchanger Network (HEN), shown in 

Appendix G. The heat exchangers from the HEN are included in the process flow diagrams 

(PFDs) in Appendix H. However, because these heat exchangers are not included in the Aspen 

simulation, the stream information to and from the pinch heat exchangers is not shown. The 

calculated minimum temperature difference across which heat can be transferred (DTmin) is 

60°F (33°C). The heating and cooling duties are satisfied through process-process interchanges 

or process-stream interchanges, thus outside utilities are not required. It is expected that any 

organization considering building a BTG plant would pursue its own heat exchanger network, 

and thus minor modifications to the heat flows within the system have been made without 

recreating the heat exchanger network. 

31
 
 

 

 

 

 

 

 

 

 

 

-1166.667 

1166.667 

2333.333 

3500.000 

4666.667 

T [°F] -1.7x108 1.7x108 3.4x108 5.1x108 6.8x108 8.5x108 

Hot Comp (Now) 

Cold Comp (Now) 

Hot Comp (Tgt) 

Cold Comp (Tgt) 

Enthalpy [Btu/h] 

Figure 7. Pinch analysis composite curves 

2.12 Water Demands 

Water is required as a reactant, a fluidizing agent, and a cooling medium in this process. As a 

reactant, it participates in reforming and water gas shift reactions. Using the BCL gasifier, it also 

acts as the fluidizing agent in the form of steam. Its cooling uses are outlined in Section 3.9. 

Water usage is becoming an increasingly important aspect of plant design, specifically with 

regard to today’s biofuel plants. Therefore, a primary design consideration for this process was 

the minimization of fresh water requirements, which therefore meant minimizing the cooling 

water demands and recycling process water as much as possible. Air-cooling was used in place 

of cooling water in several areas of the process (e.g., distillation condensers, compressor 

interstage cooling). 

Table 17 quantifies the particular water demands of this design. Roughly 36% of the fresh water 

demand is for boiler feed makeup, with most of the remainder used as cooling water makeup. 

Some of this water is directly injected into the gasifier, but other system losses (blowdown) also 

exist. This process design requires 6.5 gallons of fresh water for each gallon of gasoline 

produced. 

The option of a dry cooling tower exists. The principal advantage of the dry cooling tower is that 

it significantly reduces or even eliminates the use of water as the cooling medium in the cooling 

tower. In this case, eliminating the water usage from the cooling tower would lower the process 

water usage to 2.4 gallons of water per gallon of gasoline. However, the disadvantages of the dry 

cooling tower are the increase in capital costs and increase in power consumption (information 

on dry cooling towers can be found in the Public Interest Energy Research (PIER) Program 

Project Summary [California Energy Commission 2002]). 

32
 
 

 

 

 

 

 

 

 

 

 

Table 17. Process Water Demands 

Fresh Water Demands 

lb/h (kg/h) 

Cooling tower makeup 

Boiler feed makeup 

Sand/ash wetting 

Total 

Dry Cooling 

Tower 

0 (0) 

100,577 (45,621) 

243 (110) 

100,820 (45,731) 

Wet Cooling Tower 

174,716 (79,250) 

100,577 (45,621) 

243 (110) 

275,536 (124,981) 

Overall water demand 

(gal water / gal gasoline) 

2.4 6.5 

3 Process Economics 

The total project investment (based on total equipment cost) was developed, along with variable 

and fixed operating costs. From these costs, a discounted cash flow analysis was used to 

determine the production cost of BTG gasoline. This section describes the cost areas and 

assumptions made to complete the economic analysis. 

Each piece of equipment in these processes was sized based on the mass and energy balance data 

generated from the Aspen Plus simulation. From this, capital costs (purchase cost) were 

determined from a variety of sources, including previous studies, Questimate/Aspen IPE 

software, and engineering consultants (Nexant 2006a-d). Equipment costs from Spath et al. 

(2005) were used for the front-end sections where the process was virtually the same. The amine 

system cost was obtained from Nexant. Generalized equipment (heat exchangers, compressors, 

tanks, pumps, etc.) costs were obtained using Questimate. The fixed-bed methanol synthesis 

reactor cost was also estimated using Questimate. Distillation columns and other separation units 

were evaluated individually, and costs were scaled using values taken from the EDR (Phillips et 

al. 2007). The capital and operating costs of converting the methanol into a finished gasoline 

product were added to the detailed design. Table 18 gives installed equipment costs by plant 

area. 

Installation cost factors were used to develop a total installed cost (TIC) from the total purchased 

equipment cost. These factors are identical to those used previously (Spath et al. 2005; Phillips et 

al. 2007). The same is true for the indirect cost factors used to calculate engineering, 

construction, legal, and project contingency costs. The costs and cost factors for the current 

design are shown in Table 19 and Table 20. Costs are in 2007 dollars. This economic analysis 

does not include any royalties or license fees for use of proprietary technology. 

33
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 18. Installed Equipment Costs by Process Area 

Description of Cost Installed Cost, MM$ 

($2007) 

Feed handling & drying $25.0 

Gasification $14.6 

Tar reforming & quench $27.4 

Acid gas & sulfur removal $12.1 

Methanol synthesis - compression $10.5 

Methanol conditioning/degassing $4.8 

MTG process $21.6 

Steam system & power generation $23.1 

Cooling water & other utilities $5.9 

Total installed equipment cost $145.0 

Indirect costs $54.6 

(% of TPI) 27.4% 

Project contingency $4.2 

Total project investment (TPI) $199.6 

Table 19. General Cost Factors in Determining Total Installed Equipment Costs 

% of TPEC 

Total purchased equipment cost (TPEC) 100 

Purchased equipment installation 39 

Instrumentation and controls 26 

Piping 31 

Electrical systems 10 

Buildings (including services) 29 

Yard improvements 12 

Total installed cost (TIC) 247 

The indirect costs (non-manufacturing, fixed-capital investment costs) were also estimated using 

cost factors as per Spath et al. (2005). The factors are shown in Table 20 and have been put in as 

percentages in terms of total purchased equipment cost, total installed cost (TIC), and total 

project investment (TPI is the sum of the TIC and the total indirect costs). 

Table 20. Cost Factors for Indirect Costs 

Indirect Costs % of TPEC % of TIC % of TPI 

Engineering 32 13 9 

Construction 34 14 10 

Legal and contractors fees 23 9 7 

Project contingency 7.4 3 2 

Total indirect costs 96.4 39 28 

34
 
35 

Table 21 shows the breakdown of operating cost contribution to PGP. The first column of 

numbers is based on the total thermal energy of products. The second column is based on 

gasoline product. The last column is the percentage of the total contribution for each line. 

Table 21. Breakdown of Operating Cost Contribution to PGP

Operating Costs Cents/MMBtu Products

(Cents/GJ Products)

Cents/Gal Gasoline

(Cents/L Gasoline)

% of PGP

Feedstock 692.1 (656.0) 81.3 (21.5) 41.7%

Natural gas 0.0 (0.0) 0.0 (0.0) 0.0%

Catalysts 6.8 (6.5) 0.8 (0.2) 0.4%

Olivine 8.2 (7.8) 1.0 (0.3) 0.5%

Other raw materials 26.9 (25.5) 3.2 (0.9) 1.6%

Waste disposal 10.6 (10.0) 1.2 (0.3) 0.6%

Electricity transfer 0.0 (0.0) 0.0 (0.0) 0.0%

Electricity 0.0 (0.0) 0.0 (0.0) 0.0%

Fixed costs 243.7 (231.0) 28.6 (7.6) 14.7%

Co-product credits 0.0 (0.0) 0.0 (0.0) 0.0%

Capital depreciation 175.0 (165.9) 20.6 (5.4) 10.5%

Average income tax 126.8 (120.2) 14.9 (3.9) 7.6%

Net annual profit (after tax) 369.9 (350.6) 43.4 (11.5) 22.3%

PGP (total) 1,660.1 (1,573.5) 194.9 (51.5) 100.0%

The operating costs used in this analysis are shown in Table 22. The annualized costs of each are 

shown later. Specific catalyst compositions and costs are generally proprietary and not readily 

available. All costs have been adjusted to $2007 from their original values. The fixed operating 

costs, including labor and maintenance, are given below in Table 24 and Table 25. 

No co-product credits were included in Table 21 because co-products already are factored into 

the final PGP based on their production rate and energy content. This method gave a PGP for 

LPG of $1.53/gallon ($0.40/liter) in $2007. During 2010, the wholesale market value for propane 

varied between $1.30/gallon ($0.34/liter) and $1.47/gallon ($0.39/liter) (EIA 2010). If a $2009 

co-product credit of $1.35/gallon ($0.36/liter) is used for LPG (adjusted to $2007), the PGP for 

gasoline increases to $2.00/gallon ($0.54/liter). 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 22. Variable Operating Costs 

Variable Information and Operating Cost 

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

was sized for a gas hourly space velocity (GHSV) of 2,476/h 

based on the operation of the tar reformer at NREL’s TCPDU, 

where GHSV is measured at standard temperature and pressure. 

Initial fill, then a replacement of 0.01% per day of the total catalyst 

volume. 

Price: $6.03/lba ($13.29/kg) 

Methanol synthesis catalyst Initial fill, then replaced every 4 years based on typical catalyst 

lifetime. 

Catalyst inventory based on GHSV of 8,000/h. 

Price: $9.69b/lb ($21.36/kg) 

MTG catalyst (ZSM-5) Initial fill, then replaced every 1 year based on typical catalyst 

lifetime. 

Catalyst inventory based on 

1.84c lb fresh feed/h/lb catalyst (1.84 kg fresh feed/h/kg catalyst) 

Price: $53.40b/lb ($117.73/kg) 

Alkylation catalysts Hydrofluoric acid: 0.3 lb/bbld total alkylate (9 g/L) 

Caustic: 0.2 lb/bbld total alkylate (6 g/L) 

Price: $1,595/U.S. ton ($1,758/metric tonne) 

Solids disposal cost Price: $28.80/U.S. ton ($31.75/metric tonne) 

Diesel fuel Usage: 10 gallon/h plant-wide use (38 L/h) 

Price: $2.20/gallon ($0.58/liter) 

Chemicals Boiler chemicals – price: $2.27/lb ($5.00/kg) 

Cooling tower chemicals – price: $1.36/lb ($3.00/kg) 

LO-CAT chemicals – price: $177/metric tonne of sulfur produced 

($161/U.S. ton) 

Wastewater The wastewater is sent off-site for treatment. 

Price: $2.44/100 ft3 ($0.86/m3) 

a GAO 2005. b Jones and Zhu 2009. c Schreiner 1978. d Gary and Handwerk 1994. 

The cash flow assumptions used in this analysis are shown in Table 23. The PGP, or minimum 

product selling price necessary to achieve a 10% internal rate of return (IRR) over a 20-year 

plant life, is estimated by adjusting the value of the products to give an NPV equal to zero. The 

economic parameters chosen for a cash-flow analysis can have an enormous impact on the 

overall economics. For example, return on investment (ROI) and debt/equity financing are 

assumptions that are often debated and that vary from company to company. Although these can 

be evaluated by using sensitivity analyses, they are often so significant that they mask the PGP 

sensitivity to research-impacted parameters. 

36
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 23. Economic Parameters 

Assumption	 Value 

Internal rate of return (after-tax)	 10% 

Debt/equity	 0%/100% 

Plant life	 20 years 

General plant depreciation	 200% DDB 

General plant recovery period	 7 years 

Steam plant depreciation	 150% DDB 

Steam plant recovery period	 20 years 

Construction period	 2.5 years 

1st 6 months expenditures	 8% 

Next 12 months expenditures	 60% 

Last 12 months expenditures	 32% 

Start-up time	 6 months 

Revenues	 50% 

Variable costs	 75% 

Fixed costs	 100% 

Working capital	 5% of total capital investment 

Land	 6% of total purchased equipment cost (cost 

taken as an expense in the 1st construction year) 

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

factor was used. This also covers costs such as general plant maintenance, plant security, 

janitorial services, and communications. The 2003 PEP yearbook lists the national average 

loaded labor rate at $37.66 per hour. Using the salaries in Table 24 along with the 60% general 

overhead factor from Aden et al. (2002) gave an average loaded labor rate of $30 per hour. To 

more closely match the PEP yearbook average, the overhead factor was raised to 95% and the 

resulting loaded labor rate is $37.87 per hour ($2007) (Phillips et al. 2007). The number of plant 

personnel was adjusted to reflect the additional MTG process areas. In Table 24, these increases 

are shown in parentheses relative to the EDR values. 

37
 
 

 

 

 

 

 

 

 

 

 

 

 

Table 24. Labor Costs 

Position Salary Number Total Cost 

Plant manager $110,000 1 $110,000 

Plant engineer $65,000 2 $130,000 

Maintenance supervisor $60,000 1 $60,000 

Lab manager $50,000 1 $50,000 

Shift supervisor $45,000 5 $225,000 

Lab technician $35,000 4 $140,000 

Maintenance technician $40,000 12 (8+4) $480,000 

Shift operators $40,000 30 (20+10) $1,200,000 

Yard employees $25,000 12 $300,000 

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

Total salaries ($2002) $2,770,000 

($2007) $3,580,000 

Table 25. Other Fixed Costs 

Cost Item Factor Cost 

General overhead 95% of total salaries $3,401,000 

Maintenance 2% of total project investment $3,992,000 

Insurance & taxes 2% of total project investment $3,992,000 

4 Economics - Results 

Using the previously mentioned discounted cash flow parameters and cost information, we 

calculated a minimum gasoline selling price, or PGP. Results are given based on the specific 

product values and on a gallon ethanol equivalent (gee) basis for comparison to other processes. 

No sales prices for co-products were needed for this analysis because all co-products were 

included in the final PGP based on their production rate and energy content. 

The BTG results are shown in Table 26 along with the results from the EDR. 

38
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 26. Process and Economic Results Summary for 2012 BTG Case 

BTG Process – EDRa – 

Price calculated on As reported in reference 

energy content of all 

products 

Feedstock rate (plant size), dry metric tonne/day 

(dry U.S. ton/day) 

On-line time, h/yr 

Total yield, gallons/dry U.S. ton (liters/dry metric 

tonne) 

Fuel products, MMgal/yr (ML/year) 

Total project investment ($MM) 

PGP or minimum product selling price 

$/gal gasoline ($/L gasoline) 

$/gal ethanol equivalent ($/L ethanol 

equivalent) 

$/gal LPG ($/L LPG) 

$/MMBtu fuel, HHV basis ($/GJ fuel) 

Reference Year dollars 

Internal rate of return (IRR) 

Feedstock cost, $/dry U.S. ton ($/dry metric 

tonne) 

Equity % of total plant investment 

Carbon efficiency to desired product 

Overall plant efficiency (LHV-basis) 

2,000 (2,205) 

8,406 

55.1 (229.9) – gasoline 

9.3 (38.8) – LPG 

42.5 (160.9) – gasoline 

7.1 (26.9) – LPG 

199.6 

$1.95 ($0.52)
 

$1.39 ($0.37)
 

$1.53 ($0.40)
 

$16.60 ($15.73)


 

2007
 

10%
 

$50.70 ($55.89)
 

100% 

27.9% 

42.6% 

2,000 (2,205) 

8,406 

80.1 (334.2) – ethanol 

94.1 (392.7) – mixed 

alcohols 

61.8 (233.9) – ethanol 

72.6 (274.8) – mixed 

alcohols 

210.2 

$1.28 ($0.34)
 

2007
 

10%
 

$50.70 ($55.89)
 

100%
 

27.2%
 

47.4%
 a As published in the EDR (Phillips et al. 2007) with adjustments to U.S. $2007 and same feedstock cost of 

$50.70/dry U.S. ton ($55.89/dry metric tonne). Note this PGP is different than the one cited in the OBP MYPP, 

referenced in the Executive Summary. The difference is attributed to an adjustment made in the alcohol synthesis 

target. 

4.1 Cost Contribution for Gasoline 

The contribution of each process area to the PGP is shown graphically in Figure 8. The cost 

contributions are divided into capital, variable, and fixed operating costs. Feedstock cost is the 

largest single contributor to the PGP and represents essentially all of the variable operating cost. 

39
 
 

 

 

 

 

 

 

Capital Recovery Charge Catalysts, Raw Materials, & Waste Fixed Costs 

Feedstock 

Tar Reforming & Quench 

Methanol to Gasoline Conversion 

Feed Handling & Drying 

Alcohol Synthesis + Compression 

Cooling Water & Other Utilities 

Acid Gas & Sulfur Removal 

Gasification 

Alcohol Degassing 

Steam System & Power Generation 

Gasoline Finishing 

$0.00 $0.10 $0.20 $0.30 $0.40 $0.50 $0.60 $0.70 $0.80 $0.90 

Figure 8. Cost breakdown by area in $/gallon 

4.2 Sensitivity Analyses 

Sensitivity analyses were conducted to quantify the cost impacts that various process variables 

and different process configurations would have on the PGP. We used several methods to 

evaluate sensitivities in the process model, especially focusing on operating parameters for areas 

that are not based on commercial processes or costs that are uncertain, such as catalyst costs and 

lifetimes. A list of the variables considered is shown in Table 27. 

Each parameter was changed to a high and a low value as indicated, and the new PGP was 

calculated. The new PGP can be obtained by multiplying the percent change in PGP, in fraction 

form, by $1.95/gallon ($0.52/liter) of gasoline, and then adding or subtracting the result to or 

from $1.95/gallon ($0.52/liter) of gasoline. 

40
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

Table 27. List of Variables for Sensitivity Analyses 

Process Parameters Base Values Parameter Ranges % Change % Change in 

in PGP 

Parameter 

Feedstock cost, $/ton 50.9 (56.1) 20–85 (22–94) -61 and +66 -26 and +28 

($/tonne) 

Unconverted syngas 87 67–93 -23 and +7 +4 and 0 

recycle amount ­

methanol synthesis, % 

Other raw materials, 0.27 (0.26) 0.20–0.34 (0.19–0.33) -25 and +25 0 and 0 

$/MMBtu ($/GJ) 

Tar reformer catalyst 9,340 934–18,680 (1,030– -90 and 0 and 0 

cost, $/ton ($/tonne) (10,300) 20,590) +200 

Olivine cost, $/ton 172.9 (190.6) 17.29–345.8 (19.1–381.2) -90 and -1 and +1 

($/tonne) +200 

Internal rate of return, % 10 0–30 -100 and -27 and +91 

+200 

Total project investment, 199.6 172.1–500 -12 and -3 and +62 

MM$ +150 

Plant size, dry tonne/day 2,000 (2,205) 600–10,000 (660–11,000) -70 and +102 and -38 

(dry ton/day) +400 

Feed moisture content, 50 25–70 -50 and +40 -4 and +19 

wt % 

Installation factor 2.47 1.98–3.46 -20 and +40 -3 and +4 

H2:CO ratio 2.1 1.47–2.73 -30 and +30 0 and +1 

Methanol synthesis 8.63 (19.03) 5.25–26.25 (11.57–57.87) -39 and 0 and +1 

catalyst, $/lb ($/kg) +204 

MTG – ZSM-5 catalyst, 53.4 (117.7) 5.25–150 (11.6–330.7) -90 and -2 and +5 

$/lb ($/kg) +181 

Stream factor 0.96 0.85–1 -11 and + 4 +6 and -2 

Feedstock Cost: As the cost by area graph (Figure 8) shows, feedstock cost is a major portion of 

the overall product cost and the single largest contributor to operating costs. A 38% change in 

the feed cost (higher or lower) resulted in a 16% change in the PGP. This is illustrated along with 

other parameters’ impact in Figure 9. The more vertical a line is on the graph, the more sensitive 

the PGP is to a change in the given parameter. 

Tar Reforming Conversions: If, instead of the targets, the proven tar reformer conversions 

(Phillips et al. 2004) are input into the model, the result is a 6% increase in PGP. 

Dry Cooling Tower Scenario: If a dry cooling tower is utilized, a 4% increase in PGP is 

predicted. 

Gasoline Yields: If the gasoline yields decrease by 10% there is a 12% increase in the PGP. 

41
 
 

 

 

 

 

 

 

 

 

 

 

Fixed Bed Scenario: A simulation utilizing a fixed bed MTG reactor was used to predict a 33% 

increase in PGP with a 7.5:1 recycle ratio of gases to methanol (this scenario purchased 27 MW 

of electricity). 

Alkylation: If the alkylation unit is removed the PGP increases by 3%. 

Unconverted Syngas Recycle Ratio: To optimize the methanol synthesis process, a high (87%) 

recycle of unconverted syngas to the methanol synthesis reactor was input into the model. 

Varying the recycle down to 67% increased the PGP by 4%. 

Synthesis Reactor Pressure: The methanol synthesis catalyst and process are commercial with 

well-established operating parameters. No sensitivity analysis was done on pressure because the 

operating pressure is not an uncertainty. 

Synthesis Reactor CO Conversion: As with the reactor pressure, the performance of the methanol 

catalysts is well established. No sensitivity analysis was done on CO conversion. 

H2:CO Ratio: The H2:CO ratio was varied ±30% from the base case level of 2.1. This sensitivity 

reflects only the impact of making syngas with that ratio. It does not reflect any impacts on the 

synthesis catalyst from using a different ratio. An H2:CO ratio less than 2 most likely will have a 

negative impact on methanol production because it is below the stoichiometric ratio desired for 

making methanol from CO and H2. Operating with a higher ratio than 2.1 should not negatively 

impact the synthesis catalyst performance. The impact of varying this ratio by ±30% resulted in 

PGP changes of 0%–1%. 

Other Synthesis Parameters: Sensitivity analyses were not conducted for the following synthesis 

parameters: GHSV, temperature, pressure, reactor design, and sulfur concentration. A more 

rigorous synthesis model, based on kinetic parameters, is needed to have any confidence in a 

parameter’s impact on the methanol production and the resulting PGP for gasoline. 

Methanol Condensation Temperature: The impact of using a lower temperature in the methanol 

condensation train was evaluated. A temperature of 40°F was set for the gas/methanol separation 

vessel. As a quick estimate, no additions were made to the models to account for the additional 

equipment and energy costs for chilling the process stream. Although more methanol was 

collected at 40°F compared to the 90°F used in the base case, the PGP for gasoline was 

unchanged. Adding the equipment and energy necessary to achieve a 40°F temperature would 

result in a higher PGP. No attempt was made to adjust the plant’s energy and material integration 

to perhaps improve on the result. 

Gasoline Finishing Processes: The process used for finishing the “crude” gasoline from the 

MTG reactors is primarily distillation, which separates the various fractions and results in a 

fungible gasoline product. The model used for evaluating this part of the process has highly 

coupled unit operations. For example, a side draw is taken from the splitter column to provide 

reflux liquid for the absorber column. The liquid exiting the absorber bottom is pumped to the 

de-ethanizer column for its reflux liquid. A sizeable amount of butanes and propanes is generated 

in the process. Opportunities exist for upgrading these paraffins beyond what was done in the 

current model using available technologies. An alkylation unit was added to the model to form 

isooctane from isobutane and 2-butene, but straight-chain butanes were not upgraded. The 

42
 
 

 

 

 

 

 

 

economic impact of upgrading the butanes should be evaluated to ascertain whether the 

increased capital and operating expenses would be justified. 

Figure 9. Sensitivity analysis for biomass-to-gasoline process 

The horizontal axis in Figure 9 is the percent change in the sensitivity parameter, and the vertical 

axis is the percent change in the PGP. A more vertical line for the specified variable 

demonstrates greater volatility on the PGP when variations occur. It is clear that the factors with 

the greatest impact on the PGP are feedstock cost, feedstock moisture level, plant size, and the 

economic factors: internal rate of return and total project investment. The gasoline PGP is very 

much dependent on these factors. As is shown by the horizontal nature of the lines representing 

the catalyst costs, their contribution to the PGP is minimal, and variations in the catalyst costs 

should not greatly affect the PGP. 

5 Conclusions 

This report summarizes the results of a conceptual process design, detailed mass and energy 

balance model, and economic analysis for gasoline from biomass via gasification, methanol 

synthesis, and the MTG process. The analysis showed that gasoline could potentially be 

43
 
 

 

 

 

 

 

 

 

 

 

 

 

produced from a thermochemical biorefinery with a PGP of $1.95/gallon ($0.52/liter) (U.S. 

$2007). This has a gallon ethanol equivalent (gee) price of $1.39 ($0.37 liter ethanol equivalent). 

The PGP for gasoline on a gee basis is comparable to $1.57 for ethanol ($0.41/liter ethanol) from 

the OBP MYPP. These two values are close enough to be considered the same, given the level of 

uncertainty in the estimates used in the modeling of both processes. The benefits of one product 

over the other will likely come from site-specific, local externalities such as shipping costs, 

fungibility with existing infrastructure (e.g., pipelines), and local markets and preferences. 

The process outlined here shares common gasification and tar reforming research needs with the 

thermochemical ethanol process. The processes downstream of the gas cleaning and conditioning 

section are significantly different in terms of development. The methanol synthesis process has 

been commercial for many years and is well established. The MTG conversion process has been 

demonstrated on a large scale for the fixed bed scenario and on a pilot scale for the fluidized bed 

scenario, but it is not as widely established as the methanol process. The technology for the fixed 

bed MTG process is available as a licensed product from ExxonMobil. The gasoline process is a 

good fit for petroleum companies wanting to add biomass-derived products to their product 

portfolio. The gasoline from this process is fungible with existing refinery products and 

infrastructure. Additional analyses should investigate the economic feasibility of increasing 

alkylate production from light hydrocarbons – technologies that are well known by petroleum 

refiners and technology vendors. 

In conclusion, the results from this preliminary evaluation indicate great potential for producing 

gasoline from biomass via thermochemical biomass conversion to syngas and the MTG process, 

and thus warrant a more detailed study. Future work areas of interest include obtaining better 

process information on the MTG section of the plant, especially equipment and operating costs; 

increasing the heat integration throughout the process; scale-up of the MTG fluidized bed 

reactor; testing the MTG reactor and catalyst with methanol from biomass-derived syngas; 

testing of the MTG fluidized bed reactor at higher pressure; and evaluating the possibility of 

selling raw MTG gasoline and refining it in an existing refinery. 

6 Acknowledgments 

This work was supported by the U.S. Department of Energy’s Biomass Program, under Contract 

DE-AC36-08-GO28308 with the National Renewable Energy Laboratory. Colleagues at NREL 

have contributed to and assisted in developing many of the models and tools that were the 

starting point for this work. Major contributors to the model on which this paper is based were 

Dr. Richard Bain, Pamela Spath, Andy Aden, and John Jechura. Editing was provided by Sara 

Havig, NREL Communications Office. Many useful and insightful comments were provided by 

internal reviewers, Andy Aden and Dr. Richard Bain, and external reviewers Dr. Robert Brown 

at Iowa State University, Dr. Nicholas Petrellis, a chemical processing technologies consultant 

and recently retired from BP, Peter Tijm at PV Enterprises, Inc., and Jim Wykowski and Dave 

Payton at GEMI, University of Houston. Thank you to all contributors. 

44
 
 

 

 

 

 

 

 

 

 

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Appendix A. List of Acronyms
 

BCL Battelle Columbus Laboratory MeOH Methanol 

BFW Boiler Feed Water MMBtu Million British Thermal Units 

bpd Barrels per Day MoS2 Molybdenum Disulfide 

BTG Biomass-to-Gasoline MTBE Methyl Tertiary-Butyl Ether 

BTU British Thermal Unit MTG Methanol-to-Gasoline 

CE Chemical Engineering MW Megawatts 

Magazine 

CFM Cubic Feet per Minute MYPP Biomass Multi-Year Program Plan 

CH4 Methane MYTP Biomass Multi-Year Technical Plan 

CO Carbon Monoxide NPV Net Present Value 

CO2 Carbon Dioxide NREL National Renewable Energy Laboratory 

DCFROR Discounted Cash Flow Rate of NRTL Non-Random Two Liquid activity 

Return coefficient method 

DOE U.S. Department of Energy OBP Office of the Biomass Program 

DME Dimethylether PEP SRI Consulting Process Economics 

Program 

EDR Ethanol Design Report PFD Process Flow Diagram 

(Phillips et al. 2007) 

EIA Energy Information PGP Plant Gate Price 

Administration 

EtOH Ethanol PNNL Pacific Northwest National Laboratory 

FTL Fischer-Tropsch Liquids PPMV Parts Per Million by Volume 

FY Fiscal Year psia Pounds per Square Inch (absolute) 

GAO Government Accountability RKS- Redlich-Kwong-Soave equation of state 

Office BM with Boston-Mathius modifications 

GHSV Gas Hourly Space Velocity RVP Reid Vapor Pressure 

GJ GigaJoule SMR Steam Methane Reformer 

gpm Gallons per Minute TCPDU NREL’s Thermochemical Process 

Development Unit 

H2 Hydrogen TIC Total Installed Cost 

HF Hydrofluoric (acid) tpd Short Tons per Day 

HHV Higher Heating Value TPI Total Project Investment 

IRR Internal Rate of Return WGS Water Gas Shift 

kWh Kilowatt-hour WWT Wastewater Treatment 

LHV Lower Heating Value ZSM-5 Zeolite Catalyst 

LO-CAT Hydrogen Sulfide Removal 

Technology 

LPG Liquefied Petroleum Gas 

LTS Low Temperature Shift 

MESP Minimum Ethanol Selling Price 

50
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Appendix B. Comparison of Aspen Model to Four MTG 

Compositions from Literature 

DOE Mobil - Fluidized Bed Aspen KO - PH -

Fixed 

Bed -

300psi 

(2.1 

MPa) 

DOE -

Fixed 

Bed -

300psi 

(2.1 

Mpa) 

Mobil -

Fluidized 

Bed -

60psi 

(413 

kPa) 

Aspen 

-

200psi 

(1.4 

MPa) 

Coke 0.1 

Acetone 0.5 

Formic Acid 0.5 

Methanol 0.0 

Dimethylether 0.0 

Water 0.0 0.0 

Carbon Monoxide 0.0 

Carbon Dioxide 0.2 

Hydrogen 0.0 

Methane Methane CH4 1.0 1.0 0.8 0.8-0.9 1.0 

Ethane Ethane C2H6 0.6 0.4 0.1-0.25 0.6 

Ethene Ethylene C2H4 0.5 0.0 4.4-6.5 0.5 

Propane Propane C3H8 16.2 5.0 4.6 2.8-3.7 5.2 

Propene Propene C3H6 1.0 0.2 5.4-8.0 1.0 

N-Butane n-Butane N-C4H10 5.6 12.0 2.7 0.9-1.5 5.0 

I-Butane Isobutane I-C4H10 18.7 8.6 10.8-

15.7 

7.5 

Butenes BUTENE 1.3 1.0 1.1 1.0 

Dimethylbutanes 1.0-1.2 

N-Pentane n-Pentane PENTANE 1.3 14.0 1.4 0.3-0.5 14.5 

I-Pentanes Isopentane 7.8 12.0 10.1-

12.4 

Pentenes Pentenes PENTENE 0.5 2.0 2.2 2.2-3.1 2.1 

Cyclopentane Cyclopentane 0.2 0.1-0.4 

Methylcyclopentane Methylcyclopentane 4.3 1.3 0.6-0.7 

Methylpentanes 5.4-6.0 

N-Hexane n-Hexane HEXANE 14.0 0.7 4.1-5.4 14.5 

I-Hexanes 12.2 

51
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Hexenes Hexenes HEXENE 2.0 1.8 0.3-0.5 2.1 

Cyclohexane Cyclohexane 0.03-0.3 

Methylcyclohexane 0.4 

N-heptane C7-PON HEPTANE 6.0 0.2 3.3-4.3 6.2 

I-Heptanes 5.6 

Heptenes HEPTENE 4.0 2.0 4.2 

1,3-Dicyclopentane, 1.6 

cis 

N-Octane C8-Paraffins, Olefins, OCTANE 2.0 2.9-3.9 2.1 

Naphthenes 

I-Octanes IOCTANE 1.9 

Octenes OCTENE 5.0 2.4 5.2 

N- 2.4 

Propylcyclopentane 

N-Nonane C9-Paraffins, Olefins, NONANE 1.0 0.1 1.6-2.4 2.1 

Naphthenes 

I-Nonanes 0.8 

Nonenes NONENE 2.0 1.0 2.1 

N- 0.6 

Butylcyclopentane 

I-Decanes C10-Paraffins, 0.2 0.1-0.4 

Olefins, Naphthenes 

Decenes 0.4 

Benzene Benzene C6H6 1.7 0.2 0.0 1.8 

Toluene Toluene TOLUENE 10.5 2.0 1.8 1.4-4.0 1.9 

Ethylbenzene Ethylbenzene ETHBENZ 0.8 0.5 0.16- 1.0 

0.18 

Xylenes, meta+para Xylenes XYLENE 17.2 9.0 6.6 5.5-6.3 8.5 

O-Xylene 1.8 

1,2,4- Trimethylbenzenes 124TMB 7.5 10.0 7.0 7.2-8.4 9.5 

Trimethylbenzene 

1,3,5- 0.3 

Trimethylbenzene 

Methylethylbenzenes 1.3-1.4 

52
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

P-Ethyltoluene 2.5 

I-Propylbenzene Propylbenzenes 0.1 0.02-

0.03 

Naphthalene Naphthalenes C10H8 3.3 0.0-1.1 0.0 

1,2,4,5- 1,2,4,5- 1245TMB 8.0 4.2 1.9-3.4 0.7 

Tetramethylbenzene Tetramethylbenzene 

1,2,3,5- 1,2,3,5- 1235TMB 0.6 1.0-1.8 0.0 

Tetramethylbenzene Tetramethylbenzene 

1,2,3,4- 1,2,3,4- 0.2 0.4-0.6 0.8 

Tetramethylbenzene Tetramethylbenzene 

P-Diethylbenzene 1.9 

Other C10 Benzenes 1.0-1.4 

Penta- 0.2 0.7 

methylbenzene 

C11 Benzenes 0.3-1.7 

Other Aromatics 0.7-2.6 

2- 0.2 

Methylnaphthalene 

The compositions are given in wt %. The sources are the following: KO = Kirk-Othmer Encyclopedia of Chemical 

Technology (does not specify reactor type or pressure); PH = Probstein and Hicks (1982); DOE = 1978 DOE report 

(Schreiner 1978); Fluid Bed from Liederman et al. 1978. 

53
 
 

 

 

 

Appendix C. NREL Biorefinery Design Database Description and 

Summary 

NREL’s Process Engineering Team has developed a database of primary information on all of the 

equipment in the benchmark model. This database contains information about the cost, reference year, 

scaling factor, scaling characteristic, design information and back-up cost referencing. The information is 

stored in a secure database and can be directly linked to the economic portion of the model. In addition to 

having all of the cost information used by the model, it has the ability to store documents pertaining to the 

piece of equipment. These include sizing and costing calculations and vendor information when available. 

The following summarizes the important fields of information contained in the database. A partial listing 

of the information is attached for each piece of equipment. Additional information from the database is 

contained in the equipment cost listing in Appendix D. 

Equipment Number:AB 

Equipment Name:AB 

Associated PFD: 

Equipment Category:A 

Equipment Type:A 

Equipment Description:A 

Number Required:B
 

Number Spares:B
 

Scaling Stream:B
 

Base Cost:B
 

Cost Basis:A
 

Cost Year:B
 

Base for Scaling:B
 

Base Type:
 

Base Units:
 

Installation Factor:B
 

Installation Factor Basis:
 

Scale Factor Exponent:B
 

Scale Factor Basis:
 

Material of Construction:A
 

Notes:
 

Document:
 

Design Date:
 

Modified Date:
 

Unique identifier, the first letter indicates the equipment type and the first 

number represents the process area, e.g., P-301 is a pump in Area 300 

Descriptive name of the piece of equipment 

PFD number on which the piece of equipment appears, e.g., PFD-P800-A101 

Code indicating the general type of equipment, e.g., PUMP 

Code indicating the specific type of equipment, e.g., CENTRIFUGAL for a 

pump 

Short description of the size or characteristics of the piece of equipment, e.g., 

20 gpm, 82 ft head for a pump 

Number of duplicate pieces of equipment needed 

Number of on-line spares 

Stream number or other characteristic variable from the ASPEN model by 

which the equipment cost will be scaled 

Equipment cost 

Source of the equipment cost, e.g., ICARUS or VENDOR 

Year for which the cost estimate is based 

Value of the scaling stream or variable used to obtain the base cost of the 

equipment 

Type of variable used for scaling, e.g., FLOW, DUTY, etc. 

Units of the scaling stream or variable, e.g., KG/HR, CAL/S 

Value of the installation factor. Installed Cost = Base Cost x Installation 

Factor 

Source of the installation factor value, e.g., ICARUS, VENDOR 

Value of the exponential scaling equation 

Source of the scaling exponent value, e.g., GARRETT, VENDOR 

Material of Construction 

Any other important information about the design or cost 

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. 

Original date for the design of this piece of equipment 

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 this Appendix.
 B These fields are part of the equipment cost listing in Appendix D.
 

54
 
 

 

 

 

EQUIPMENT_NUM 

BER EQUIPMENT_NAME EQUIPMENT_ TYPE 

MATERIAL_ 

CONST COST_ BASIS 

PFD-850-A101 

C-101 Hopper Feeder VIBRATING-FEEDER CS LITERATURE 

C-102 Screener Feeder Conveyor BELT CS LITERATURE 

C-103 Radial Stacker Conveyor BELT CS LITERATURE 

M-101 Hydraulic Truck Dump with Scale TRUCK-SCALE LITERATURE 

M-102 Hammermill CS LITERATURE 

M-103 Front End Loaders LOADER CS LITERATURE 

S-101 Magnetic Head Pulley MAGNET CS LITERATURE 

S-102 Screener SCREEN CS LITERATURE 

T-101 Dump Hopper LIVE-BTM-BIN CS LITERATURE 

T-102 Hammermill Surge Bin LIVE-BTM-BIN CS LITERATURE 

T-103 Dryer Feed Bin LIVE-BTM-BIN CS LITERATURE 

PFD-850-A102 

C-104 Dryer Feed Screw Conveyor SCREW CS LITERATURE 

C-105 Gasifier Feed Screw Conveyor SCREW 316SS LITERATURE 

K-101 Flue Gas Blower CENTRIFUGAL SS304 LITERATURE 

M-104 Rotary Biomass Dryer ROTARY-DRUM CS LITERATURE 

S-103 Dryer Air Cyclone GAS CYCLONE CS LITERATURE 

S-104 Dryer Air Baghouse Filter FABRIC-FILTER LITERATURE 

T-104 Dried Biomass Hopper VERTICAL-VESSEL CS LITERATURE 

PFD-850-A201 

C-201 Sand/ash Conditioner/Conveyor SCREW CS LITERATURE 

H-209A Pinch HX System DHE FIXED T S IPE 

H-209B Pinch HX System DHE U TUBE IPE 

H-209C Pinch HX System DHE U TUBE IPE 

K-202 Combustion Air Blower CENTRIFUGAL CS LITERATURE 

M-201 Sand/ash Cooler MISCELLANEOUS LITERATURE 

R-201 Indirectly-heated Biomass Gasifier VERTICAL-VESSEL 

CS 

w/refractory LITERATURE 

R-202 Char Combustor VERTICAL-VESSEL 

CS 

w/refractory LITERATURE 

S-201 Primary Gasifier Cyclone GAS CYCLONE 

CS 

w/refractory LITERATURE 

S-202 Secondary Gasifier Cyclone GAS CYCLONE 

CS 

w/refractory LITERATURE 

S-203 Primary Combustor Cyclone GAS CYCLONE 

CS 

w/refractory LITERATURE 

S-204 Secondary Combustor Cyclone GAS CYCLONE 

CS 

w/refractory LITERATURE 

S-205 Electrostatic Precipitator MISCELLANEOUS CS LITERATURE 

T-201 Sand/ash Bin FLAT-BTM-STORAGE CS LITERATURE 

55
 
 

 

PFD-P850-A301 

H-301A Post-Reformer Cooler #1 SHELL-TUBE IPE 

H-315D Pinch HX System EHE JACKETED IPE 

H-315DB Pinch HX System DHE FIXED T S IPE 

H-330A Pinch HX System DHE FIXED T S IPE 

H-330B Pinch HX System DHE FIXED T S IPE 

H-330C Pinch HX System DHE FIXED T S IPE 

H-330D Pinch HX System DHE FIXED T S IPE 

K-305 Regenerator Combustion Air Blower CENTRIFUGAL SS304 QUESTIMATE 

R-301A Tar Reformer Catalyst Regenerator VERTICAL-VESSEL 

CS 

w/refractory LITERATURE 

R-303 Tar Reformer VERTICAL-VESSEL 

CS 

w/refractory LITERATURE 

S-306 Tar Reformer Cyclone GAS CYCLONE CS LITERATURE 

S-307 Catalyst Regenerator Cyclone GAS CYCLONE CS LITERATURE 

PFD-P850-A302 

H-301B Post-Reformer Cooler #2 

AIR-COOLED 

EXCHANGER A214 IPE 

M-301 Syngas Quench Chamber CS LITERATURE 

M-302 Syngas Venturi Scrubber CS LITERATURE 

P-301 Sludge Pump CENTRIFUGAL CS QUESTIMATE 

P-302 Quench Water Recirculation Pump CENTRIFUGAL CS LITERATURE 

T-301 Sludge Settling Tank CLARIFIER SS304 QUESTIMATE 

T-302 Quench Water Recirculation Tank HORIZONTAL-VESSEL CS LITERATURE 

PFD-P850-A303 

H-302 Syngas Compressor Intercoolers 

AIR-COOLED 

EXCHANGER CS ICARUS 

H-303 Water-cooled Aftercooler SHELL-TUBE 

SS304CS/A2 

14 QUESTIMATE 

K-301 Syngas Compressor CENTRIFUGAL A285C QUESTIMATE 

S-301 Pre-compressor Knock-out KNOCK-OUT DRUM CS QUESTIMATE 

S-302 Syngas Compressor Interstage Knock-outs KNOCK-OUT DRUM CS ICARUS 

S-303 Post-compressor Knock-out KNOCK-OUT DRUM CS QUESTIMATE 

PFD-P850-A304 

S-310 L.P. Amine System ABSORBER OTHER 

PFD-P850-A305 

H-304 LO-CAT Preheater SHELL-TUBE 

A285C/CA44 

3 QUESTIMATE 

H-305 LO-CAT Absorbent Solution Cooler SHELL-TUBE 304SS VENDOR 

K-302 LO-CAT Feed Air Blower CENTRIFUGAL CS VENDOR 

M-303 LO-CAT Venturi Precontactor 304SS VENDOR 

M-304 LO-CAT Liquid-filled Absorber ABSORBER 304SS VENDOR 

56
 
 

 

 

 

 

 

 

 

 

 

 

 

 

P-303 LO-CAT Absorbent Solution Circulating Pump CENTRIFUGAL 304SS VENDOR 

R-304 LO-CAT Oxidizer Vessel VERTICAL-VESSEL 304SS VENDOR 

PFD-P850-A401 

R-434 Low Temperature Shift Reactor VERTICAL-VESSEL 

CS 

w/refractory QUESTIMATE 

PFD-P850-A402 

H-301C 

Reformed Syngas cooler / Synthesis Reactor 

Preheat #1 SHELL-TUBE A214 QUESTIMATE 

H-416A Pinch HX System DHE U TUBE IPE 

H-416B Pinch HX System DHE U TUBE IPE 

K-412 Purge Gas Expander CENTRIFUGAL A285C QUESTIMATE 

K-414 Unreacted Syngas Recycle Compressor CENTRIFUGAL A285C QUESTIMATE 

S-414 Methanol Flash Drum HORIZONTAL-VESSEL CS QUESTIMATE 

PFD-P850-A403 

H-410A Pinch HX System DHE U TUBE IPE 

H-410C Pinch HX System DHE U TUBE IPE 

H-410D Pinch HX System DHE U TUBE IPE 

H-410E Pinch HX System DHE U TUBE IPE 

H-410F Pinch HX System DHE U TUBE IPE 

H-410G Pinch HX System DHE U TUBE IPE 

H-411 post Reactor Syngas cooling #2 SHELL-TUBE 

A285C/CA44 

3 IPE 

H-411B post Reactor Syngas cooling #3 - Air Cooled 

AIR-COOLED 

EXCHANGER A214 QUESTIMATE 

H-413 post Reactor Syngas cooling #3 - Air Cooled SHELL-TUBE CS/A214 QUESTIMATE 

H-414 Methanol Condenser SHELL-TUBE A214 QUESTIMATE 

R-490 High pressure Synthesis Reactor VERTICAL-VESSEL 

CS 

w/refractory QUESTIMATE 

PFD-P850-A502 

D-500 Methanol Degassing Column DISTILLATION SS305 ICARUS 

H-500R Methanol Column Reboiler SHELL-TUBE SS304;CS ICARUS 

H-504B Cooler SHELL-TUBE A214 QUESTIMATE 

H-505A Pinch HX System DHE U TUBE IPE 

H-505B Pinch HX System DHE U TUBE IPE 

H-505C Pinch HX System DHE U TUBE IPE 

H-505D Pinch HX System DHE U TUBE IPE 

H-593 METHANOL Product Finishing cooler SHELL-TUBE CS LITERATURE 

K-501 Methanol Gas Expander CENTRIFUGAL A285C QUESTIMATE 

P-500B Deaerator Feed Pump CENTRIFUGAL CS QUESTIMATE 

P-500R Deaerator Feed Pump CENTRIFUGAL CS QUESTIMATE 

P-592 Methanol Product Pump CENTRIFUGAL CS ICARUS 

57
 
 

 

 

 

 

 

 

 

 

PFD-P850-A601 

H-605A Pinch HX System DHE FIXED T S IPE 

H-605B Pinch HX System DHE FIXED T S IPE 

H-605C Pinch HX System DHE FIXED T S IPE 

H-605D Pinch HX System DHE U TUBE IPE 

H-605E Pinch HX System DHE U TUBE IPE 

H-605F Pinch HX System DHE U TUBE IPE 

H-630A Pinch HX System DHE FIXED T S IPE 

H-630B Pinch HX System DHE U TUBE IPE 

H-642 Pinch HX System EHE JACKETED IPE 

M-601 Hot Process Water Softener System PACKAGE RICHARDSON 

P-603 Deaerator Feed Pump CENTRIFUGAL CS QUESTIMATE 

P-604 Boiler Feed Water Pump CENTRIFUGAL CS QUESTIMATE 

T-601 Condensate Collection Tank HORIZONTAL-VESSEL CS QUESTIMATE 

T-602 Condensate Surge Drum HORIZONTAL-VESSEL CS QUESTIMATE 

T-603 Deaerator HORIZONTAL-VESSEL CS;SS316 VENDOR 

PFD-P850-A602 

H-200A Pinch HX System EHE JACKETED IPE 

H-200B Pinch HX System DHE PRE ENGR IPE 

H-601 Steam Turbine Condenser SHELL-TUBE 

ADEN, ET. AL. 

2002 

H-620 Pinch HX System DHE FIXED T S IPE 

M-602 Extraction Steam Turbine/Generator STEAM-TURBINE VENDOR 

P-601 Collection Pump CENTRIFUGAL CS QUESTIMATE 

P-602 Condensate Pump CENTRIFUGAL SS304 QUESTIMATE 

PFD-P850-A603 

H-603 Blowdown Water-cooled Cooler SHELL-TUBE A214 QUESTIMATE 

H-607 Pinch HX System DHE FIXED T S IPE 

S-601 Blowdown Flash Drum HORIZONTAL-VESSEL CS QUESTIMATE 

T-604 Steam Drum HORIZONTAL-VESSEL CS ICARUS 

PFD-P850-A701 

K-701 Plant Air Compressor RECIPROCATING CS ICARUS 

M-701 Cooling Tower System INDUCED-DRAFT 

FIBERGLAS 

S DELTA-T98 

P-701 Cooling Water Pump CENTRIFUGAL CS QUESTIMATE 

S-701 Instrument Air Dryer PACKAGE CS RICHARDSON 

T-701 Plant Air Receiver HORIZONTAL-VESSEL CS ICARUS 

58
 
 

 

 

 

PFD-P850-A702 

M-702 Hydraulic Truck Dump with Scale TRUCK-SCALE CS VENDOR 

P-702 Firewater Pump CENTRIFUGAL CS ICARUS 

P-703 Diesel Pump CENTRIFUGAL CS ICARUS 

P-704 Ammonia Pump CENTRIFUGAL CS ICARUS 

P-705 Hydrazine Pump CENTRIFUGAL CS DELTA-T98 

T-702 Firewater Storage Tank FLAT-BTM-STORAGE A285C ICARUS 

T-703 Diesel Storage Tank FLAT-BTM-STORAGE A285C ICARUS 

T-704 Ammonia Storage Tank 

HORIZONTALSTORAGE A515 ICARUS 

T-705 Olivine Lock Hopper VERTICAL-VESSEL CS DELTA-T98 

T-706 MgO Lock Hopper VERTICAL-VESSEL CS DELTA-T98 

T-707 Hydrazine Storage Tank VERTICAL-VESSEL SS316 ICARUS 

PFD-P850-A1401 

H-1401A Pinch HX System DHE FIXED T S IPE 

H-1401B Pinch HX System DHE FIXED T S IPE 

H-1401C Pinch HX System DHE FIXED T S IPE 

H-1410 Pinch HX System DHE FIXED T S IPE 

P-1000 Methanol Intermediate Pump CENTRIFUGAL CS ICARUS 

R-1410 MTG Reactor VERTICAL-VESSEL 

T-592 Methanol Product Storage Tank FLAT-BTM-STORAGE A285C ICARUS 

PFD-P850-A1402 

H-315A Pinch HX System EHE JACKETED IPE 

H-1411B Pinch HX System DHE FIXED T S IPE 

H-1413 post MTG Reactor #3 Air Cooled SHELL-TUBE 

A285C/CA44 

5 QUESTIMATE 

H-1414 post MTG Reactor #4 water cooled SHELL-TUBE 

A285C/CA44 

6 QUESTIMATE 

H-1416 Pinch HX System DHE FIXED T S IPE 

S-1401 Mixed HC Condensation Knock-out KNOCK-OUT DRUM A-515 QUESTIMATE 

PFD-P850-A1501 

D-1502 Absorber Column DISTILLATION IPE 

D-1503 Deethanizer Column DISTILLATION IPE 

H-1503R De-ethanizer Column Reboiler SHELL-TUBE SS304;CS ICARUS 

H-1506A Pinch HX System DHE FIXED T S IPE 

H-1506B Pinch HX System DHE FIXED T S IPE 

H-1506C Pinch HX System DHE FIXED T S IPE 

H-1506D Pinch HX System DHE FIXED T S IPE 

H-1508 D1503 Reflux Heater SHELL-TUBE A214 QUESTIMATE 

H-1512B Lean Oil Recycle Cooler #2 SHELL-TUBE 

SS304CS/A2 

14 QUESTIMATE 

P-1503 Crude Hydrocarbons Pump CENTRIFUGAL CS ICARUS 

59
 
P-1503B D1503 Reboiler Pump CENTRIFUGAL CS ICARUS 

P-1508 D1503 Reflux Pump CENTRIFUGAL CS ICARUS 

PFD-P850-A1502 

D-1504 Stabilizer Column DISTILLATION IPE 

D-1505 Splitter Column DISTILLATION IPE 

D-1506 LPG/ Alkylate Splitter DISTILLATION IPE 

H-1504R Stabilizer Column Reboiler SHELL-TUBE SS304;CS ICARUS 

H-1504C Stabilizer Column Condenser SHELL-TUBE A214 QUESTIMATE 

H-1505R Splitter Column Reboiler SHELL-TUBE SS304;CS ICARUS 

H-1505 Pinch HX System DHE FIXED T S IPE 

H-1505B Pinch HX System DHE FIXED T S IPE 

P-1504B D1504 Reboiler Pump CENTRIFUGAL CS ICARUS 

P-1505B D1505 Reboiler Pump CENTRIFUGAL CS ICARUS 

R-1505 HF Alkylation Unit 

CS 

w/refractory ICARUS 

PFD-P850-A1503 

H-1591 LPG product cooler - cw SHELL-TUBE A214 IPE 

H-1593 Gasoline product cooler - cw SHELL-TUBE A214 IPE 

P-1590 LPG Product Pump CENTRIFUGAL CS ICARUS 

P-1592 Gasoline Product Pump CENTRIFUGAL CS ICARUS 

T-1590 LPG Product Storage Tank FLAT-BTM-STORAGE A285C ICARUS 

T-1592 Gasoline Product Storage Tank FLAT-BTM-STORAGE A285C ICARUS 

60
 
 

 

 

 

 

 

 

 

 

 

 

Appendix D. Individual Equipment Cost Summary
 

Total 

Equipment 

Number 

Number 

Required 

Number 

Spares Equipment Name 

Size 

Ratio 

Original Equip 

Cost (per unit) 

Base 

Year 

Original 

Equip Cost 

(Req'd & 

Spare) in 

Base Year 

Scaling 

Exponent 

Scaled 

Cost in 

Base Year 

Installation 

Factor 

Installed 

Cost in Base 

Year 

Installed 

Cost in 

$2007 

C-101 4 Hopper Feeder 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

Screener Feeder 

C-102 2 Conveyor 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

C-103 2 Radial Stacker Conveyor 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

C-104 2 

Dryer Feed Screw 

Conveyor 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

Gasifier Feed Screw 

C-105 2 Conveyor 0.98 $0 2002 $0 0.75 $0 2.47 $0 $0 

K-101 2 Flue Gas Blower 1.60 $0 2002 $0 0.75 $0 2.47 $0 $0 

M-101 4 

Hydraulic Truck Dump 

with Scale 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

M-102 2 Hammermill 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

M-103 3 Front End Loaders 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

M-104 2 Rotary Biomass Dryer 1.00 $3,813,728 2002 $7,627,455 0.75 $7,627,450 2.47 $18,839,801 $25,021,313 

S-101 2 Magnetic Head Pulley 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

S-102 2 Screener 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

S-103 2 Dryer Air Cyclone 1.60 $0 2002 $0 0.75 $0 2.47 $0 $0 

S-104 2 Dryer Air Baghouse Filter 0.98 $0 2002 $0 0.75 $0 2.47 $0 $0 

T-101 4 Dump Hopper 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

T-102 1 Hammermill Surge Bin 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

T-103 2 Dryer Feed Bin 1.00 $0 2002 $0 0.75 $0 2.47 $0 $0 

T-104 2 Dried Biomass Hopper 0.98 $0 2002 $0 0.75 $0 2.47 $0 $0 

61
 
 

 

 

 

 

 

 

 

 

A100 Subtotal $7,627,455 $7,627,450 2.47 $18,839,801 $25,021,313 

C-201 1 

Sand/ash 

Conditioner/Conveyor 0.33 $0 2002 $0 0.65 $0 2.47 $0 $0 

K-202 2 Combustion Air Blower 1.01 $0 2002 $0 0.65 $0 2.47 $0 $0 

M-201 2 Sand/ash Cooler 0.33 $0 2002 $0 0.65 $0 2.47 $0 $0 

R-201 2 

Indirectly-heated Biomass 

Gasifier 1.01 $2,212,201 2002 $4,424,402 0.65 $4,451,012 2.47 $10,994,000 $14,601,232 

R-202 2 Char Combustor 1.01 $0 2002 $0 0.65 $0 2.47 $0 $0 

S-201 2 Primary Gasifier Cyclone 1.01 $0 2002 $0 0.65 $0 2.47 $0 $0 

S-202 2 

Secondary Gasifier 

Cyclone 1.00 $0 2002 $0 0.65 $0 2.47 $0 $0 

S-203 2 

Primary Combustor 

Cyclone 1.01 $0 2002 $0 0.65 $0 2.47 $0 $0 

S-204 2 

Secondary Combustor 

Cyclone 0.99 $0 2002 $0 0.65 $0 2.47 $0 $0 

S-205 2 Electrostatic Precipitator 1.00 $0 2002 $0 0.65 $0 2.47 $0 $0 

T-201 1 Sand/ash Bin 0.33 $0 2002 $0 0.65 $0 2.47 $0 $0 

H-200A 1 Pinch HX System $10,600 2007 $10,600 0.6 $10,600 2.47 $26,182 $26,182 

H-200B 1 Pinch HX System $1,800 2007 $1,800 0.6 $1,800 2.47 $4,446 $4,446 

H-209A 1 Pinch HX System $30,600 2007 $30,600 0.6 $30,600 2.47 $75,582 $75,582 

H-209B 1 Pinch HX System $38,200 2007 $38,200 0.6 $38,200 2.47 $94,354 $94,354 

H-209C 1 Pinch HX System $54,600 2007 $54,600 0.6 $54,600 2.47 $134,862 $134,862 

A200 Subtotal $4,424,402 $4,451,012 2.47 $10,994,000 $14,601,232 

H-301A 1 Pinch HX System $98,200 2007 $98,200 0.6 $98,200 2.47 $242,554 $242,554 

H-301B 1 Post-Reformer Cooler #2 0.90 $50,300 2007 $50,300 0.6 $47,357 2.47 $116,973 $116,973 

H-301C 1 Pinch HX System $22,400 2007 $22,400 0.6 $22,400 2.47 $55,328 $55,328 

H-302 5 

Syngas Compressor 

Intercoolers 0.84 $0 2002 $0 0.65 $0 2.47 $0 $0 

H-303 1 Water-cooled Aftercooler 0.00 $20,889 2002 $20,889 0.44 $0 2.47 $0 $0 

H-304 1 LO-CAT Preheater 0.18 $4,743 2002 $4,743 0.6 $1,677 2.47 $4,142 $5,500 

62 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

H-305 1 

LO-CAT Absorbent 

Solution Cooler 0.34 $0 2002 $0 0.44 $0 2.47 $0 $0 

H-315A 

H-315D 

H-315DB 

H-330A 

H-330B 

H-330C 

H-330D 

K-301 

1 

1 

1 

1 

1 

1 

1 

1 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Syngas Compressor 0.92 

$5,400 

$83,700 

$17,600 

$18,900 

$17,500 

$35,800 

$40,700 

$3,896,834 

2007 

2007 

2007 

2007 

2007 

2007 

2007 

2002 

$5,400 

$83,700 

$17,600 

$18,900 

$17,500 

$35,800 

$40,700 

$3,896,834 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.8 

$5,400 

$83,700 

$17,600 

$18,900 

$17,500 

$35,800 

$40,700 

$3,652,085 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

$13,338 

$206,739 

$43,472 

$46,683 

$43,225 

$88,426 

$100,529 

$9,020,650 

$13,338 

$206,739 

$43,472 

$46,683 

$43,225 

$88,426 

$100,529 

$11,980,408 

K-302 1 LO-CAT Feed Air Blower 0.78 $0 2002 $0 0.65 $0 2.47 $0 $0 

K-305 1 

Regenerator Combustion 

Air Blower 1.03 $35,020 2002 $35,020 0.59 $35,651 2.47 $88,058 $116,951 

M-301 1 Syngas Quench Chamber 0.86 $0 2002 $0 0.65 $0 2.47 $0 $0 

M-302 1 Syngas Venturi Scrubber 0.86 $0 2002 $0 0.65 $0 2.47 $0 $0 

M-303 1 

LO-CAT Venturi 

Precontactor 0.78 $0 2002 $0 0.65 $0 2.47 $0 $0 

M-304 1 

LO-CAT Liquid-filled 

Absorber 0.34 $0 2002 $0 0.65 $0 2.47 $0 $0 

P-301 1 1 Sludge Pump 0.08 $3,911 2002 $7,822 0.33 $3,435 2.47 $8,484 $11,268 

P-302 1 1 

Quench Water 

Recirculation Pump 0.27 $0 2002 $0 0.65 $0 2.47 $0 $0 

P-303 1 1 

LO-CAT Absorbent 

Solution Circulating Pump 0.86 $0 2002 $0 0.65 $0 2.47 $0 $0 

R-301A 

R-303 

1 

1 

Tar Reformer Catalyst 

Regenerator 

Tar Reformer 

0.89 

0.86 

$2,429,379 

$2,212,201 

2002 

2002 

$2,429,379 

$2,212,201 

0.65 

0.65 

$2,253,216 

$2,009,880 

2.47 

2.47 

$5,565,444 

$4,964,405 

$7,391,518 

$6,593,272 

R-304 1 LO-CAT Oxidizer Vessel 0.78 $1,000,000 2002 $1,000,000 0.65 $851,685 2.47 $2,103,663 $2,793,894 

S-301 1 

Pre-compressor Knockout 0.92 $157,277 2002 $157,277 0.6 $149,809 2.47 $370,027 $491,437 

S-302 4 

Syngas Compressor 

Interstage Knock-outs 0.92 $0 2002 $0 0.6 $0 2.47 $0 $0 

63 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Post-compressor KnockS-303 1 out 0.00 $40,244 2002 $40,244 0.6 $0 2.47 $0 $0 

S-306 1 Tar Reformer Cyclone 0.86 $0 2002 $0 0.65 $0 2.47 $0 $0 

Catalyst Regenerator 

S-307 1 Cyclone 0.89 $0 2002 $0 0.65 $0 2.47 $0 $0 

S-310 1 L.P. Amine System 0.77 $3,485,685 2002 $3,485,685 0.75 $2,866,983 2.47 $7,081,449 $9,404,938 

T-301 1 Sludge Settling Tank 0.49 $11,677 2002 $11,677 0.6 $7,573 2.47 $18,706 $24,844 

Quench Water 

T-302 1 Recirculation Tank 0.86 $0 2002 $0 0.65 $0 2.47 $0 $0 

A300 Subtotal $13,594,071 $12,121,353 2.47 $29,939,741 $39,528,742 

H-410A 1 Pinch HX System $25,000 2007 $25,000 0.6 $25,000 2.47 $61,750 $61,750 

H-410C 1 Pinch HX System $123,700 2007 $123,700 0.6 $123,700 2.47 $305,539 $305,539 

H-410D 1 Pinch HX System $28,300 2007 $28,300 0.6 $28,300 2.47 $69,901 $69,901 

H-410E 1 Pinch HX System $85,100 2007 $85,100 0.6 $85,100 2.47 $210,197 $210,197 

H-410G 1 Pinch HX System $16,200 2007 $16,200 0.6 $16,200 2.47 $40,014 $40,014 

H-410F 1 Pinch HX System $21,800 2007 $21,800 0.6 $21,800 2.47 $53,846 $53,846 

H-411 1 Pinch HX System $212,000 2007 $212,000 0.6 $212,000 2.47 $523,640 $523,640 

H-411B 1 

post Reactor Syngas 

cooling #3 - Air Cooled 0.17 $388,064 2002 $388,064 0.6 $135,869 2.47 $335,597 $445,709 

H-413 1 

post Reactor Syngas 

cooling #3 - Air Cooled 0.01 $71,389 2002 $71,389 0.44 $8,338 2.47 $20,595 $27,352 

H-414 1 Methanol Condenser 0.14 $338,016 2002 $338,016 0.44 $140,431 2.47 $346,866 $460,676 

H-416A 1 Pinch HX System $16,200 2007 $16,200 0.6 $16,200 2.47 $40,014 $40,014 

H-416B 1 Pinch HX System $17,900 2007 $17,900 0.6 $17,900 2.47 $44,213 $44,213 

K-412 1 Purge Gas Expander 1.72 $642,014 2002 $642,014 0.8 $991,770 2.47 $2,449,673 $3,253,433 

K-414 1 

Unreacted Syngas 

Recycle Compressor 3.34 $403,122 2002 $403,122 0.8 $1,057,398 2.47 $2,611,772 $3,468,718 

R-434 1 

Low Temperature Shift 

Reactor 0.00 $4,965,833 2002 $4,965,833 0.7 $16,442 2.47 $40,612 $53,937 

R-490 1 

High pressure Synthesis 

Reactor 0.17 $2,026,515 2002 $2,026,515 0.56 $743,768 2.47 $1,837,106 $2,439,877 

S-414 1 Methanol Flash Drum 29.95 $14,977 2002 $14,977 0.6 $115,147 2.47 $284,412 

64
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

S-471 1 

Methanol Condensation 

Knock-out 2.04 $55,447 2002 $55,447 0.6 $84,962 2.47 $209,856 $278,712 

A400 

D-500 

H-500R 

H-504B 

H-592 

H-505A 

H-505B 

H-505C 

H-505D 

H-592B 

H-593 

K-501 

P-500B 

P-500R 

P-592 

T-592 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

Methanol Degassing 

Column 

Methanol Column 

Reboiler 

Cooler 

METHANOL Product 

Cooler 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

METHANOL Product 

Finishing cooler 

Methanol Gas Expander 

D-500 Bottoms Pump 

D-500 Reflux Pump 

Methanol Product Pump 

Methanol Product 

Storage Tank 

0.11 

1.27 

0.16 

1.00 

1.00 

1.00 

1.00 

1.00 

1.00 

1.63 

1.27 

0.01 

0.01 

2.07 

2.07 

$478,100 

$29,600 

$338,016 

$3,043 

$20,300 

$54,300 

$61,600 

$27,800 

$20,300 

$0 

$642,014 

$8,679 

$8,679 

$7,500 

$165,800 

Subtotal 

1998 

1997 

2002 

2002 

2007 

2007 

2007 

2007 

2007 

2002 

2002 

2002 

2002 

1997 

1997 

$8,939,477 

$478,100 

$29,600 

$338,016 

$3,043 

$20,300 

$54,300 

$61,600 

$27,800 

$20,300 

$0 

$642,014 

$17,358 

$17,358 

$7,500 

$165,800 

0.68 

0.68 

0.44 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.44 

0.8 

0.33 

0.33 

0.79 

0.51 

$3,328,225 

$104,412 

$34,861 

$151,190 

$3,043 

$20,300 

$54,300 

$61,600 

$27,800 

$20,300 

$0 

$775,537 

$3,036 

$3,428 

$13,348 

$240,556 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

$8,220,716 

$257,898 

$86,107 

$373,439 

$7,516 

$50,141 

$134,121 

$152,152 

$68,666 

$50,141 

$0 

$1,915,575 

$7,499 

$8,467 

$32,971 

$594,172 

$10,512,643 

$347,881 

$117,052 

$495,968 

$9,982 

$50,141 

$134,121 

$152,152 

$68,666 

$50,141 

$0 

$2,544,093 

$9,959 

$11,246 

$44,820 

$807,705 

A500 

H-601 

H-603 

H-605A 

H-605B 

H-605C 

1 

1 

1 

1 

1 

Steam Turbine 

Condenser 

Blowdown Water-cooled 

Cooler 

Pinch HX System 

Pinch HX System 

Pinch HX System 

1.39 

0.00 

$0 

$16,143 

$18,200 

$379,900 

$36,300 

65 

Subtotal 

2002 

2002 

2007 

2007 

2007 

$1,883,089 

$0 

$16,143 

$18,200 

$379,900 

$36,300 

0.71 

0.44 

0.6 

0.6 

0.6 

$1,513,711 

$0 

$0 

$18,200 

$379,900 

$36,300 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

$3,738,866 

$0 

$0 

$44,954 

$938,353 

$89,661 

$4,843,928 

$0 

$0 

$44,954 

$938,353 

$89,661 
 

 

 

 

 

 

 

 

 

H-605D 1 Pinch HX System $31,100 2007 $31,100 0.6 $31,100 2.47 $76,817 $76,817 

H-605E 1 Pinch HX System $76,400 2007 $76,400 0.6 $76,400 2.47 $188,708 $188,708 

H-605F 1 Pinch HX System $25,400 2007 $25,400 0.6 $25,400 2.47 $62,738 $62,738 

H-607 1 Pinch HX System $36,400 2007 $36,400 0.6 $36,400 2.47 $89,908 $89,908 

H-620 1 Pinch HX System $20,400 2007 $20,400 0.6 $20,400 2.47 $50,388 $50,388 

H-630A 1 Pinch HX System $28,200 2007 $28,200 0.6 $28,200 2.47 $69,654 $69,654 

H-630B 1 Pinch HX System $16,200 2007 $16,200 0.6 $16,200 2.47 $40,014 $40,014 

H-642 1 Pinch HX System $10,900 2007 $10,900 0.6 $10,900 2.47 $26,923 $26,923 

Hot Process Water 

M-601 1 Softener System 1.20 $1,031,023 1999 $1,031,023 0.82 $1,197,991 2.47 $2,959,037 $3,980,230 

Extraction Steam 

M-602 1 Turbine/Generator 1.20 $4,045,870 2002 $4,045,870 0.71 $4,599,565 2.47 $11,360,926 $15,088,550 

M-603 1 Startup Boiler 1.00 $198,351 2002 $198,351 0.6 $198,351 2.47 $489,927 $650,676 

P-601 1 1 Collection Pump 0.39 $7,015 2002 $14,030 0.33 $10,310 2.47 $25,465 $33,821 

P-602 1 1 Condensate Pump 1.39 $5,437 2002 $10,874 0.33 $12,116 2.47 $29,927 $39,747 

P-603 1 1 Deaerator Feed Pump 1.20 $8,679 2002 $17,358 0.33 $18,439 2.47 $45,544 $60,487 

P-604 1 1 Boiler Feed Water Pump 1.20 $95,660 2002 $191,320 0.33 $203,233 2.47 $501,985 $666,691 

Condensate Collection 

T-601 1 Tank 1.20 $24,493 2002 $24,493 0.6 $27,336 2.47 $67,520 $89,675 

T-602 1 Condensate Surge Drum 1.20 $28,572 2002 $28,572 0.6 $31,889 2.47 $78,765 $104,609 

T-603 1 Deaerator 1.20 $130,721 2002 $130,721 0.72 $149,135 2.47 $368,365 $489,228 

T-604 1 Steam Drum 1.20 $9,200 1997 $9,200 0.72 $10,478 2.47 $25,881 $35,182 

S-601 1 Blowdown Flash Drum 1.20 $14,977 2002 $14,977 0.6 $16,692 2.47 $41,229 $54,757 

A600 Subtotal $6,412,332 $7,154,934 2.47 $17,672,688 $22,971,769 

K-701 2 1 Plant Air Compressor 1.00 $32,376 2002 $97,129 0.34 $97,129 2.47 $239,908 $318,624 

M-701 1 Cooling Tower System 0.00 $267,316 2002 $267,316 0.78 $0 2.47 $0 $0 

M-702 1 

Hydraulic Truck Dump 

with Scale 1.00 $80,000 1998 $80,000 0.6 $80,000 2.47 $197,600 $266,544 

M-703 1 Flue Gas Stack 0.32 $51,581 2002 $51,581 1 $16,705 2.47 $41,262 $54,801 

P-701 1 1 Cooling Water Pump 1.33 $158,540 2002 $317,080 0.33 $348,290 2.47 $860,276 $1,142,541 

P-702 1 1 Firewater Pump 1.00 $18,400 1997 $36,800 0.79 $36,800 2.47 $90,896 $123,562 

P-703 1 1 Diesel Pump 1.00 $6,100 1997 $12,200 0.79 $12,200 2.47 $30,134 $40,963 

66 
 

 

P-704 

P-705 

S-701 

T-701 

T-702 

T-703 

T-704 

T-705 

T-706 

T-707 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

Ammonia Pump 

Hydrazine Pump 

Instrument Air Dryer 

Plant Air Receiver 

Firewater Storage Tank 

Diesel Storage Tank 

Ammonia Storage Tank 

Olivine Lock Hopper 

MgO Lock Hopper 

Hydrazine Storage Tank 

1.00 

1.00 

1.00 

1.00 

1.00 

1.00 

1.00 

1.00 

1.00 

1.00 

$5,000 

$5,500 

$8,349 

$7,003 

$166,100 

$14,400 

$287,300 

$0 

$0 

$12,400 

1997 

1997 

2002 

2002 

1997 

1997 

1997 

1998 

1998 

1997 

$10,000 

$5,500 

$16,698 

$7,003 

$166,100 

$14,400 

$287,300 

$0 

$0 

$12,400 

0.79 

0.79 

0.6 

0.72 

0.51 

0.51 

0.72 

0.71 

0.71 

0.93 

$10,000 

$5,500 

$16,698 

$7,003 

$166,100 

$14,400 

$287,300 

$0 

$0 

$12,400 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

$24,700 

$13,585 

$41,244 

$17,297 

$410,267 

$35,568 

$709,631 

$0 

$0 

$30,628 

$33,577 

$18,467 

$54,777 

$22,973 

$557,708 

$48,350 

$964,657 

$0 

$0 

$41,635 

A700 

P-1000 

H-1401A 

H-1401B 

H-1401C 

H-1410 

H-1411B 

H-1412 

H-1413 

H-1414 

H-1416 

R-1410 

S-1401 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 

1 Methanol Feed Pump 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

Pinch HX System 

post MTG Reactor #1 

post MTG Reactor #3 Air 

Cooled 

post MTG Reactor #4 

water cooled 

Pinch HX System 

MTG Reactor 

Mixed HC Condensation 

Knock-out 

2.07 

0.47 

0.37 

0.94 

0.25 

0.27 

$7,500 

$61,300 

$21,900 

$32,600 

$17,600 

$121,000 

$26,400 

$26,700 

$4,743 

$0 

$3,700,000 

$55,447 

Subtotal 

1997 

2007 

2007 

2007 

2007 

2007 

2007 

2007 

2002 

2007 

1999 

2002 

$1,381,507 

$15,000 

$61,300 

$21,900 

$32,600 

$17,600 

$121,000 

$26,400 

$26,700 

$4,743 

$0 

$3,700,000 

$55,447 

0.79 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.6 

0.65 

0.6 

$1,110,525 

$26,697 

$61,300 

$21,900 

$32,600 

$17,600 

$121,000 

$16,741 

$14,715 

$4,564 

$0 

$1,502,667 

$25,514 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

2.47 

$2,742,997 

$65,941 

$151,411 

$54,093 

$80,522 

$43,472 

$298,870 

$41,351 

$36,346 

$11,272 

$0 

$3,711,587 

$63,020 

$3,647,545 

$89,639 

$151,411 

$54,093 

$80,522 

$43,472 

$298,870 

$41,351 

$36,346 

$14,971 

$0 

$4,992,494 

$83,698 

A1400 

D-1502 

D-1503 

1 

1 

Absorber Column 

Deethanizer Column 

1.04 

0.86 

$347,300 

$236,400 

67 

Subtotal 

2007 

2007 

$4,082,690 

$347,300 

$236,400 

0.68 

0.68 

$1,845,298 

$357,838 

$214,037 

2.47 

2.47 

2.47 

$4,557,886 

$883,859 

$528,672 

$5,886,866 

$883,859 

$528,672 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

D-1504 1 Stabilizer Column 0.87 $236,400 2007 $236,400 0.68 $214,421 2.47 $529,619 $529,619 

D-1505 1 Splitter Column 0.81 $347,300 2007 $347,300 0.68 $301,841 2.47 $745,548 $745,548 

D-1506 1 LPG/ Alkylate Splitter 1.13 $347,300 2007 $347,300 0.68 $377,954 2.47 $933,547 $933,547 

De-ethanizer Column 

H-1503R 1 Reboiler 0.22 $29,600 1997 $29,600 0.68 $10,676 2.47 $26,369 $35,846 

Stabilizer Column 

H-1504C 1 Condenser 0.13 $338,016 2002 $338,016 0.44 $139,141 2.47 $343,677 $456,441 

Stabilizer Column 

H-1504R 1 Reboiler 0.74 $29,600 1997 $29,600 0.68 $24,071 2.47 $59,454 $80,821 

H-1505 1 Pinch HX System $18,000 2007 $18,000 0.6 $18,000 2.47 $44,460 $44,460 

H-1505B 1 Pinch HX System $21,100 2007 $21,100 0.6 $21,100 2.47 $52,117 $52,117 H-1505C 1 

Spliter Column 

Condenser 0.03 $338,016 2002 $338,016 0.44 $71,403 2.47 $176,364 $234,231 

H-1505R 1 Splitter Column Reboiler 0.15 $29,600 1997 $29,600 0.68 $8,247 2.47 $20,370 $27,690 

H-1506A 1 Pinch HX System $18,900 2007 $18,900 0.6 $18,900 2.47 $46,683 $46,683 

H-1506B 1 Pinch HX System $17,600 2007 $17,600 0.6 $17,600 2.47 $43,472 $43,472 

H-1506C 1 Pinch HX System $17,600 2007 $17,600 0.6 $17,600 2.47 $43,472 $43,472 

H-1506D 1 Pinch HX System $21,600 2007 $21,600 0.6 $21,600 2.47 $53,352 $53,352 

H-1508 1 Pinch HX System $18,100 2007 $18,100 0.6 $18,100 2.47 $44,707 $44,707 

H-1512B 1 

Lean Oil Recycle Cooler 

#2 0.06 $20,889 2002 $20,889 0.44 $6,032 2.47 $14,899 $19,788 

H-1591 1 LPB product cooler - cw 0.00 $90,000 2007 $90,000 0.44 $0 2.47 $0 $0 

H-1593 1 

Gasoline product cooler -

cw 12.29 $30,500 2007 $30,500 0.44 $91,982 2.47 $227,196 $227,196 

P-1503 1 

Crude Hydrocarbons 

Pump 0.90 $7,500 1997 $7,500 0.79 $6,915 2.47 $17,079 $23,217 

P-1503B 1 D1503 Reboiler Pump 0.91 $7,500 1997 $7,500 0.79 $6,987 2.47 $17,259 $23,461 

P-1504B 1 D1504 Reboiler Pump 0.80 $7,500 1997 $7,500 0.79 $6,258 2.47 $15,457 $21,011 

P-1505B 1 D1505 Reboiler Pump 0.39 $7,500 1997 $7,500 0.79 $3,577 2.47 $8,836 $12,011 

P-1505C 1 D1505 Reflux Pump 0.32 $7,500 1997 $7,500 0.79 $3,071 2.47 $7,585 $10,310 

P-1508 1 D1503 Reflux Pump 0.12 $7,500 1997 $7,500 0.79 $1,423 2.47 $3,515 $4,778 

P-1590 1 LPG Product Pump 0.09 $7,500 1997 $7,500 0.79 $1,114 2.47 $2,751 $3,739 

P-1592 1 Gasoline Product Pump 0.74 $7,500 1997 $7,500 0.79 $5,940 2.47 $14,671 $19,944 

R-1506 1 Refrigeration for 

LPG/Alkylate Splitter 

0.02 $700,879 2007 $700,879 0.6 $56,448 1.6 $90,317 $90,317 

68 
 

 

 

R-1590 

R-1505 

T-1590 

T-1592 

1 

1 

1 

1 

Refrigeration for 

liquefying LPG 

HF Alkylation Unit 

LPG Product Storage 

Tank 

Gasoline Product Storage 

Tank 

0.02 

0.25 

0.09 

0.74 

$700,879 

$7,080,000 

$165,800 

$165,800 

2007 

1999 

1997 

1997 

$700,879 

$7,080,000 

$165,800 

$165,800 

0.6 

0.65 

0.51 

0.51 

$57,859 

$2,875,306 

$48,401 

$142,625 

1.6 

2.47 

2.47 

2.47 

$92,575 

$7,102,005 

$119,551 

$352,285 

$92,575 

$9,552,978 

$162,515 

$478,888 

A1500 Subtotal $10,096,221 $5,112,591 2.47 $12,628,100 $15,493,644 

Equipment Cost $58,441,244 $44,265,099 2.47 $109,334,794 $142,507,681 

69
 
70 

Appendix E. Economic Summary Page from Excel Spreadsheet 

Energy Plant Gate Price ($/MMBtu) $16.60

Gasoline Plant Gate Price ($/gal) $1.95 $1.39/gee

LPG Plant Gate Price* ($/gal) $1.53

Electricity Plant Gate Price ($/kW-hr) $0.0567

Gasoline Production at Operating Capacity (MM Gal / year) 42.5

Gasoline Product Yield (gal / Dry US Ton Feedstock) 55.1

LPG Production at Operating Capacity (MM Gal / year) 7.1

LPG Product Yield (gal / Dry US Ton Feedstock) 9.3

Delivered Feedstock Cost $/Dry US Ton $51

Internal Rate of Return (After-Tax) 10%

Equity Percent of Total Investment 100%

Capital Costs Operating Costs (cents/mmBtu product) cents/gal gasoline

 Feed Handling & Drying $25,000,000 Feedstock 692.2 81.3 41.7%

 Gasification $14,600,000 Natural Gas 0.0 0.0 0.0%

 Tar Reforming, Quench, & Compression $27,400,000 Catalysts 6.8 0.8 0.4%

 Acid Gas & Sulfur Removal $12,100,000 Olivine 8.2 1.0 0.5%

 Alcohol Synthesis - Compression $10,400,000 Other Raw Materials 26.9 3.2 1.6%

 Alcohol Degassing $4,800,000 Waste Disposal 10.6 1.2 0.6%

 MTG Process $21,600,000 Electricity Transfer 0.0 0.0 0.0%

 Steam System & Power Generation $23,100,000 Electricity 0.0 0.0 0.0%

 Cooling Water & Other Utilities $5,900,000 Fixed Costs 243.8 28.6 14.7%

Total Installed Equipment Cost $144,900,000 Co-product credits 0.0 0.0 0.0%

Capital Depreciation 175.0 20.6 10.5%

Indirect Costs 54,600,000 Average Income Tax 126.8 14.9 7.6%

 (% of TPI) 27.4% Average Return on Investment 369.9 43.4 22.3%

 Project Contingency 4,200,000 PGP (Total) 1660.3 195.0 100.0%

Operating Costs ($/yr)

Total Project Investment (TPI) $199,500,000 Feedstock $39,100,000

Natural Gas $0

Installed Equipment Cost per Annual Gallon $3.27 Catalysts $200,000

Total Project Investment per Annual Gallon $4.50 Olivine $500,000

Other Raw Matl. Costs $600,000

Loan Rate N/A Waste Disposal $600,000

Term (years) N/A Electricity Transfer Charge $0

Capital Charge Factor 0.190 Electricity $0

Fixed Costs $13,800,000

Co-product credits @ $0.00 per gal $0

Capital Depreciation $9,900,000

Average Income Tax $7,200,000

Gasifier Efficiency - HHV % 75.3 Average Return on Investment $20,900,000

Gasifier Efficiency - LHV % 74.9

Efficiency to Gasoline - HHV % 37.6 Total Plant Electricity Usage (KW) 34,314

Efficiency to Gasoline - LHV % 37.7 Electricity Produced Onsite (KW) 34,320

Overall Plant Efficiency - HHV % 42.6 Electricity from OBL (KW) 0

Overall Plant Efficiency - LHV % 42.6 Electricity Purchased from Grid (KW) 0

 Electricity Sold to Grid (KW) 6

Plant Hours per year 8406

% 96.0% Steam Plant + Turboexpander Power Generated (hp) 48,415

 Used for Main Compressors (hp) 0

 Used for Electricity Generation (hp) 48,415

Plant Electricity Use (KWh/gal product) 9.1

Gasification & Reforming Steam Use (lb/gal) 14.4

Water use (gal water/gal gasoline) 6.5

Specific Operating Conditions

Feed rate 2,000 dry tonnes/day

2,205 dry tons/day

Feedstock Cost $50.70 $/dry ton

$51.17 $/maf ton

*Cost referenced to 60°F; composition is 27wt% C3's and 73wt%C4's, 4ppmwH2O.

Gasoline via Biomass Gasification/MeOH Synthesis and MTG 

Process Engineering Analysis

2,000 Dry Metric Tonnes Biomass per Day

All Values in 2007$

BCL Gasifier, Tar Reformer, Sulfur Removal, WGS & Cu/ZnO/Al2O3 & ZSM-5 catalysts, Fuel Purification, Steam-Power Cycle

2012 Market Target Case: 2010 Tar Reforming Goal & MTG Production
 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

Appendix F. Discounted Cash Flow Rate of Return Summary
 

DCFROR Worksheet 

Year -2 -1 0 1 2 3 4 5 6 7 

Fixed Capital Investment $19,220,429 $118,664,036 $63,287,486 

Working Capital $9,888,670 

Loan Payment $0 $0 $0 $0 $0 $0 $0 

Loan Interest Payment $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

Loan Principal $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

Fuel Sales $70,559,608 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 

By-Product Credit $14,607 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 

Total Annual Sales $70,574,215 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 

Annual Manufacturing Cost 

Raw Materials $34,255,310 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 

Tar reforming catalysts $722,179 

Steam reforming catalysts $0 $0 

ZnO $0 $0 

Methanol catalysts $515,085 $0 $0 $0 $515,085 $0 $0 

MTG - ZSM5 Catalyst $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 

Baghouse Bags $466,183 $466,183 

Other Variable Costs $2,573,587 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 

Fixed Operating Costs $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 

Total Product Cost $54,804,297 $58,344,206 $58,344,206 $58,344,206 $58,859,291 $58,810,389 $58,344,206 

71
 
DCFROR Worksheet continued 

8 9 10 11 12 13 14 15 16 17 18 19 20 

($9,888,670) 

$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

$94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 $94,079,478 

$19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 $19,476 

$94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 $94,098,954 

$39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 $39,148,925 

$0 $0 

$0 $0 

$0 $515,085 $0 $0 $0 $515,085 $0 $0 $0 $515,085 $0 $0 $0 

$2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 $2,485,503 

$466,183 $466,183 

$2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 $2,923,327 

$13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 $13,786,450 

$58,344,206 $58,859,291 $58,344,206 $58,810,389 $58,344,206 $58,859,291 $58,344,206 $58,344,206 $58,810,389 $58,859,291 $58,344,206 $58,344,206 $58,344,206 

72
 
 

 

 

 

 

 

 

 

 

 

 

 

 

 

DCFROR Worksheet continued 

Year -2 -1 0 1 2 3 4 5 6 

Annual Depreciation 

General Plant 

DDB $47,221,405 $33,729,575 $24,092,554 $17,208,967 $12,292,119 $8,780,085 

SL $23,610,702 $19,675,585 $16,864,787 $15,057,846 $14,340,806 $14,340,806 

Remaining Value $118,053,512 $84,323,937 $60,231,384 $43,022,417 $30,730,298 $21,950,213 

Actual $47,221,405 $33,729,575 $24,092,554 $17,208,967 $14,340,806 $14,340,806 

Steam Plant 

DDB $2,437,386 $2,254,582 $2,085,488 $1,929,077 $1,784,396 $1,650,566 

SL $1,624,924 $1,582,163 $1,544,806 $1,513,001 $1,486,997 $1,467,170 

Remaining Value $30,061,091 $27,806,509 $25,721,021 $23,791,944 $22,007,549 $20,356,982 

Actual $2,437,386 $2,254,582 $2,085,488 $1,929,077 $1,784,396 $1,650,566 

Net Revenue ($33,888,872) ($229,409) $9,576,707 $16,616,705 $19,114,462 $19,297,194 

Losses Forward ($33,888,872) ($34,118,280) ($24,541,574) ($7,924,869) $0 

Taxable Income ($33,888,872) ($34,118,280) ($24,541,574) ($7,924,869) $11,189,593 $19,297,194 

Income Tax $0 $0 $0 $0 $4,363,941 $7,525,905 

Annual Cash Income $15,769,919 $35,754,748 $35,754,748 $35,754,748 $30,875,722 $27,762,660 

Discount Factor 1.21 1.1 1 0.909090909 0.826446281 0.751314801 0.683013455 0.620921323 0.56447393 

Annual Present Value $225,493,427 $14,336,290 $29,549,379 $26,863,072 $24,420,974 $19,171,394 $15,671,298 

Total Capital Investment + Interest $23,256,719.02 $130,530,439.89 $73,176,155.70 

73
 
DCFROR Worksheet continued 

7 8 9 10 11 12 13 14 15 16 17 18 19 20 

$6,271,489 

$14,340,806 

$15,678,723 

$14,340,806 

$1,526,774 $1,412,266 $1,306,346 $1,208,370 $1,117,742 $1,033,911 $956,368 $884,640 $818,292 $756,920 $700,151 $647,640 $599,067 $554,137 

$1,454,070 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 

$18,830,209 $17,417,943 $16,111,597 $14,903,228 $13,785,485 $12,751,574 $11,795,206 $10,910,566 $10,092,273 $9,335,353 $8,635,201 $7,987,561 $7,388,494 $6,834,357 

$1,526,774 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 $1,448,478 

$19,887,169 $34,306,271 $33,791,185 $34,306,271 $33,840,088 $34,306,271 $33,791,185 $34,306,271 $34,306,271 $33,840,088 $33,791,185 $34,306,271 $34,306,271 $34,306,271 

$0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 $0 

$19,887,169 $34,306,271 $33,791,185 $34,306,271 $33,840,088 $34,306,271 $33,791,185 $34,306,271 $34,306,271 $33,840,088 $33,791,185 $34,306,271 $34,306,271 $34,306,271 

$7,755,996 $13,379,446 $13,178,562 $13,379,446 $13,197,634 $13,379,446 $13,178,562 $13,379,446 $13,379,446 $13,197,634 $13,178,562 $13,379,446 $13,379,446 $13,379,446 

$27,998,752 $22,375,303 $22,061,101 $22,375,303 $22,090,931 $22,375,303 $22,061,101 $22,375,303 $22,375,303 $22,090,931 $22,061,101 $22,375,303 $22,375,303 $22,375,303 

0.513158118 0.46650738 0.424097618 0.38554329 0.3504939 0.318630818 0.28966438 0.26333125 0.23939205 0.217629136 0.19784467 0.17985879 0.163507991 0.148643628 

$14,367,787 $10,438,244 $9,356,060 $8,626,648 $7,742,737 $7,129,461 $6,390,315 $5,892,117 $5,356,470 $4,807,630 $4,364,671 $4,024,395 $3,658,541 $3,325,946 

($1,469,887.74) 

74
 
Appendix G. Heat Exchanger Network 

*Note: It is expected that anyone building this type of BTG plant would create their own 

heat exchanger network, and thus due to time restrictions, the heat exchanger network has 

not been updated to reflect recent model modifications. 

Name Duty Areas (ft

2

) Shls/Ft Spec Stream 

Name 

Tin (oF) Tout 

(

o

F) 

1 H-200A (N1) 3.85E+6 74.47 1 Duty 286 1627.19 1600.20 

74.45 1.000 200 382.75 494.46 

2 H-200B (N2) 191000.0 6.57 1 Duty 311 1158.48 1157.00 

6.57 1.000 200 494.46 500.00 

3 H-209A (N3) 9.84E+6 1446.53 1 Duty 504 417.34 361.92 

1398.51 0.967 209 184.55 273.89 

4 H-209B (N4) 2.17E+6 1647.78 1 Duty 412 320.80 309.35 

1581.82 0.960 209 273.89 293.58 

5 H-209C (N5) 1.17E+7 2926.27 1 Duty 412 482.00 420.12 

2588.45 0.885 209 293.58 400.00 

6 H-315D (N6) 1.37E+7 615.22 1 Duty 286 1816.31 1719.89 

569.53 0.926 315D 533.40 1624.00 

7 H-416B (N7) 1.32E+7 332.69 1 Duty 286 1719.89 1627.19 

325.80 0.979 416 140.77 1199.76 

8 H-410F (N8) 3.20E+7 675.29 1 Duty 286 1600.20 1375.78 

671.94 0.995 410 328.46 483.45 

9 H-605F (N9) 3.70E+7 1032.60 1 Duty 286 1375.78 1116.21 

1027.00 0.995 605 382.75 466.36 

10 H-642 (N10) 2.31E+6 81.59 1 Duty 286 1116.21 1100.00 

81.59 1.000 642 464.99 465.00 

11 H-301A (N11) 1.14E+8 6067.67 2 Duty 301 1621.87 560.71 

6066.54 1.000 642 464.27 464.98 

12 H-605E (N12) 1.99E+7 5538.15 2 Duty 301 560.71 370.80 

5251.13 0.948 605 337.75 382.75 

13 H-1505A (N13) 642000.0 69.79 1 Duty 301 370.80 364.69 

69.64 0.998 R1505 111.30 200.00 

14 H-410A (N14) 5.48E+6 700.95 1 Duty 301 364.69 312.44 

690.27 0.985 410 130.00 186.48 

15 H-301C (N15) 4.30E+6 669.40 1 Duty 301 312.44 271.47 

667.25 0.997 620 140.00 150.08 

16 H-505A (N16) 2.57E+6 496.49 1 Duty 301 271.47 247.00 

493.32 0.994 505 129.97 151.84 

17 H-416A (N17) 3.82E+6 685.66 1 Duty 301 247.00 241.07 

683.55 0.997 416 90.91 140.77 

18 Cooling Tower 4.31E+6 Duty 301 241.07 200.00 

(N18) 

19 H-607 (N19) 1.08E+8 1438.08 1 Duty 311 2860.00 2003.65 

1407.88 0.979 607 462.23 900.00 

20 H-1416 (N20) 235000.0 3.49 1 Duty 311 2003.65 2002.35 

3.49 1.000 1416 90.00 798.85 

21 H-505D (N21) 5.00E+7 875.38 1 Duty 311 2002.35 1609.24 

861.14 0.984 505 273.89 700.00 

22 H-1401C (N22) 5.26E+7 1295.25 1 Duty 311 1609.24 1196.22 

1248.30 0.964 1401 210.78 680.00 

75
 
23 H-315A (N23) 1.39E+6 39.94 1 Duty 311 1196.22 1185.32 

39.89 0.999 315A 202.00 570.00 

24 H-410G (N24) 3.42E+6 114.22 1 Duty 311 1185.32 1158.48 

114.21 1.000 410 483.45 500.00 

25 H-315DB(N52) 2.80E+6 100.75 1 Duty 311 1157.00 1152.00 

100.74 1.000 315D 513.00 533.40 

26 H-330A (N25) 169000.0 53.10 1 Duty 592 205.34 203.30 

53.09 1.000 330 130.92 133.11 

27 H-330B (N26) 580000.0 203.73 1 Duty 504 203.27 200.00 

203.53 0.999 330 133.11 140.66 

28 H-330C (N27) 8.96E+6 2007.24 1 Duty 504 338.49 288.02 

1832.78 0.913 330 140.66 257.25 

29 H-330D (N28) 1.10E+7 2491.88 1 Duty 504 479.13 417.34 

2174.54 0.873 330 257.25 400.00 

30 H-630B (N29) 252000.0 121.61 1 Duty 630 234.37 234.00 

121.60 1.000 410 186.48 187.75 

31 H-410C (N30) 3.20E+6 2022.92 2 Duty 412 278.36 224.52 

1864.98 0.922 410 187.75 237.10 

32 H-410D (N31) 4.16E+6 924.03 1 Duty 504 361.92 338.49 

917.15 0.993 410 237.10 257.25 

33 H-410E (N32) 1.47E+7 5292.15 2 Duty 412 398.41 320.80 

5005.81 0.946 410 257.25 328.46 

34 H-411 (N33) 2.06E+7 12855.29 3 Duty 411 224.52 200.00 

12647.44 0.984 620 150.08 198.00 

35 H-605D (N34) 4.11E+6 1235.68 1 Duty 412 420.12 398.41 

1228.44 0.994 605 328.46 337.75 

36 H-505C (N35) 5.07E+6 3058.17 1 Duty 412 309.35 278.36 

2600.20 0.850 505 223.89 273.89 

37 H-505B (N36) 8.46E+6 2852.38 1 Duty 504 288.02 240.38 

2539.01 0.890 505 151.84 223.89 

38 H-1506D (N37) 1.48E+6 517.25 1 Duty 504 240.38 232.05 

505.28 0.977 1506 129.97 200.00 

39 H-1401A (N38) 5.11E+6 3165.76 1 Duty 504 232.05 203.27 

2733.82 0.864 1401 151.84 197.46 

40 H-1508 (N39) 165000.0 48.93 1 Duty 592 251.85 249.85 

48.76 0.997 1508 141.01 200.00 

41 Cooling Tower 3.67E+6 Duty 592 249.85 205.34 

(N40) 

42 H-1506B (N41) 272000.0 66.31 1 Duty 592 203.30 200.00 

66.25 0.999 1506 101.89 114.75 

43 H-605A (N42) 707000.0 165.73 1 Duty R1500 370.80 300.00 

165.39 0.998 605 233.43 235.03 

44 H-605B (N43) 3.71E+7 22318.80 5 Duty 1411 354.57 277.31 

21666.52 0.971 605 235.03 318.88 

45 H-605C (N44) 4.24E+6 1425.73 1 Duty 1412 449.70 354.57 

1376.62 0.966 605 318.88 328.46 

46 H-630A (N45) 2.59E+6 1374.95 1 Duty 630 255.59 234.37 

1357.87 0.988 620 198.00 204.48 

47 H-1505B (N46) 783000.0 456.80 1 Duty 1505 252.88 236.89 

455.32 0.997 620 204.48 206.32 

48 H-1411B (N47) 1.31E+7 7244.42 2 Duty 1411 277.31 250.00 

7097.65 0.980 620 206.32 237.10 

49 H-1410 (N48) 1.32E+6 99.85 1 Duty 1410 778.91 752.00 

99.85 1.000 642 464.98 464.99 

50 H-1401B (N49) 1.49E+6 541.17 1 Duty 1505 283.36 252.88 

531.99 0.983 1401 197.46 210.78 

76
 
51 H-1506A (N50) 227000.0 32.33 1 Duty 1512 319.68 205.84 

32.07 0.992 1506 91.15 101.89 

52 H-1506C (N51) 322000.0 70.63 1 Duty 1592 257.46 200.00 

69.68 0.987 1506 114.75 129.97 

77
 
Appendix H. Process Flow Diagrams
 

78
 
79
80
81
82
83
84
85
86
87
88
89
90
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92
93
94
95
96
97
98
99
100
 

 

 

 

 

 

 

 

 

 

Appendix I. Comparison to other Biomass-to-Gasoline and 

Methanol-to-Gasoline Published Costs 

I.1 Comparison to “Techno-economic Analysis for the Conversion of 

Lignocellulosic Biomass to Gasoline via the Methanol-to-Gasoline (MTG) 

Process” by S. B. Jones and Y. Zhu from Pacific Northwest National Laboratory 

(PNNL) 

The report “Techno-economic Analysis for the Conversion of Lignocellulosic Biomass to 

Gasoline via the Methanol-to-Gasoline (MTG) Process” by S. B. Jones and Y. Zhu from Pacific 

Northwest National Laboratory (PNNL) predicts a plant gate price (PGP) that is approximately 

65% higher than the PGP predicted in the NREL BTG study. A principal difference between the 

two studies is the state of the various technologies throughout the process. PNNL’s report 

utilizes and references proven states of technology. The report by NREL analyzes future states of 

the technologies and predicts the potential of the process. For the gasification and syngas cleanup 

sections in the NREL report, the 2012 targets used were set by DOE and are defined in the OBP 

MYPP (Multi-Year Program Plan). In the MTG section of the NREL report, a fluidized bed 

reactor was utilized (proven at the pilot scale) instead of the fixed bed reactor utilized by PNNL 

(proven at the commercial scale). The states of technologies have a large impact on the capital 

and operating costs required for the process. Other differences include but are not limited to 

stream factor, year dollars used (2007 vs. 2008), and feedstock cost. 

In the table below, the PNNL assumptions are imposed on the NREL model cumulatively and 

the price per gallon gasoline calculated. 

Table I-1. Comparison of NREL and PNNL Model Assumptions and Effect on PGP 

Assumptions NREL Price PNNL Price 

(NREL→PNNL) per Gallon per Gallon 

Original Cases 

LPG Contribution by Energy 

Value or Co-Product Credit 

Years Dollars Used 

Feedstock Cost 

Stream Factor 

Total Project Investment 

Electricity Purchase 

Other 

Energy Value→Co-Product Credit 

2007→2008

$50.70→$60.00

0.96→0.90

$213M→$336Ma 

0→1.52 kWh/gallon 

1.95 

1.96 

2.06 

2.21 

2.29 

3.03 

3.14 

3.20 

3.20
 

a 54% of the difference in Total Project Investment is attributed to the front end of the process (including feed 

handling, drying, gasification, tar reforming, quench, acid gas removal, and syngas compression) while 46% is 

attributed to the back end of the process (including methanol synthesis, MTG process, and gasoline refining). 

Explanations and other differences: 

1.	 Some of the equipment cost differences in the syngas cleanup section are due to an 

additional steam reformer used by PNNL; this steam reformer was removed from NREL 

analysis designs when tar reformer targets were met. 

2.	 Equipment cost differences for the methanol-to-gasoline process include differences in 

the types of MTG reactors chosen. NREL selected a fluidized bed reactor, while PNNL 

101
 
 

 

 

 

 

 

 

 

 

 

selected a fixed bed reactor. The fixed bed case requires a 7.5:1 gas recycle loop to the 

reactor for heat management purposes. This very large recycle results in overall higher 

costs. 

3.	 The tar reformer conversions in the 2012 State of Technology targets set by the U.S. 

Department of Energy represent a prediction of tar reforming capabilities in the near 

future. The tar reformer conversions used in the PNNL report are cited from Spath et al. 

(2005) and are more representative of demonstrated technology in 2005. Improvements in 

tar/methane reforming technology are currently being developed and demonstrated at 

NREL. 

4.	 Other differences not listed in the table above include: 

A. The NREL report includes an alkylation unit to convert 2-butene and isobutane 

into isooctane, resulting in 2% higher gasoline yields. 

B. The NREL report uses a dedicated water gas shift reactor to obtain the correct 

H2:CO ratio for the methanol reactor. 

C. The NREL report includes greater methanol recovery after the methanol reactor 

by using a product recovery system that includes a multiple flash drum 

configuration. 

D. The process in the NREL report produces all of the necessary electricity for the 

process. The PNNL design requires more than 7 MW of electricity to be 

purchased from the grid, accounting for $0.13 of PNNL’s minimum fuel selling 

price (MFSP). 

I.2 Comparison of Total Project Investment to Published Cost Information from 

the New Zealand MTG Commercial Plant (Seddon 2006) 

A report by Seddon (2006) gives some financial insight into the New Zealand MTG commercial 

plant. It states that the total cost was $1,475 million in 1985 U.S. dollars ($2,900 million in 2007 

U.S. dollars). It also states that 40% of this cost was directly from interest and inflation and thus 

it is assumed that the total project investment (TPI) is $1,740 million in 2007 U.S. dollars. The 

New Zealand plant (200 million gallons per year, Seddon 2006) is approximately 4.5 times larger 

than the BTG plant (43 million gallons per year). Other cost increasing factors include: 

•	 It was built on an active volcano and thus required extensive engineering and expensive 

infrastructure mitigation for seismic activity. 

•	 It was a first of a kind plant, and a first of a kind plant typically costs double what an nth 

plant costs. 

102
 
 

 

 

 

 

Appendix J. External Reviewer Comments and Responses 

Thank you to our external reviewers: Peter Tijm, Dr. Robert Brown, Dr. Nicholas Petrellis, Jim 

Wykowski, and Dave Payton. We worked to address all comments and suggestions as described 

in the ‘Response’ column. 

General Comments: 

• Stated that yields and efficiencies were consistent with what he had expected. 

• Thought the work looked thorough and well done. 

• Asked what percentage of capital costs were for the syngas, and was pleased with the response 

of 76%. 

• Questioned what temperature was in the low temperature shift reactor and if the steam entering 

is superheated or not. He was satisfied with response of 750°F and the steam is superheated. 

• Reviewed overall report and focused on MTG part as requested. 

• Report uses same feedstock, nth plant, detailed design, and economic analysis approaches as 

other NREL reports. 

• Using nth plant basis reportedly helps in guiding research. However doesn’t tend to highlight 

uncertainties and barriers to commercial development and readiness. Estimated cost 

uncertainties for individual technologies make comparisons between technologies problematic 

unless the differences are overwhelming. 

Suggestion/Comment Response 

Thought the process would most likely need a guard 

bed ahead of the tar reformer containing the nickelbased catalyst. 

A sensitivity was performed finding that there 

was no change in the PGP if a guard bed was 

added to the process. 

Suggested that the economics could possibly be 

improved by selling the crude gasoline, which then 

could be sent to a refinery (potential to save capital). 

Added to ‘Conclusion’ as a ‘Future work area 

of interest.’ 

Estimated that feedstock actually costs 

approximately $75/ton. 

The feedstock cost is set in the OBP MYPP and 

remains $50.70/ton. However, a sensitivity is 

included for feedstock cost in Section 4.2. 

Noted that he saw room for refinement opportunities 

in the thermal analysis. 

Added to ‘Conclusion’ as a ‘Future work area 

of interest.’ 

Suggested a 10% IRR is too low for companies to 

invest. 

10% IRR is standard in NREL Design Reports 

and in order to maintain consistency will be 

kept at 10%. However, there is a sensitivity 

included in Section 4.2 addressing Internal Rate 

of Return. 

Suggested including information of recent MTG 

processes (DKRW – Wyoming and Synthesis 

Energy Systems). 

Information included in Section 1.1. 

103
 
 

 

 

 

 

 

 

 

Suggested noting butane and 1,2,4,5­

tetramethylbenzene differences in Appendix B. 

Appendix B was revisited and corrected. An 

error was caught that the refined gasoline 

stream from the Aspen simulation was listed 

when the crude gasoline composition should 

have been listed. 

Suggested incorporating some information on Exxon 

Mobil’s licensing fees for the MTG process. 

Exxon Mobil’s licensing department was 

contacted. However, they were unable to 

release any cost-related information. 

Suggested decreasing the range of olivine and tar 

reformer costs in the sensitivity chart. 

Ranges were decreased. 

Suggested updating the acronym list. List was updated. 

Suggested including a PGP or price difference if 

methane conversion was the proven level of 50%. 

The PGP of $2.05 including current tar 

reforming capabilities was included in the 

report in Section 2.3. 

Noted PFD a402 was used when a403 should have 

been referenced. 

Correction was made. 

Suggested adding capacity to beginning of Executive 

Summary. 

Added capacity to Executive Summary. 

Noted that economies of scale are not always 

applicable, especially for very large scale equipment. 

Sizing factors were not changed in order to 

maintain consistency with NREL Design 

Reports. 

Suggested diesel price was far too low. Increased diesel fuel cost to $63,626 cents/ton 

($2.20/gallon). 

Suggested clarifying ton vs. tonne. Ton is now specified as U.S. Ton and Tonne is 

now specified as Metric Tonne. 

Questioned if the foundation of the work in the 

Ethanol Design Report was targeting Ethanol or 

Mixed Alcohols. 

It is now specified in Section 1.1 that the work 

done for the Ethanol Design Report was for 

ethanol via a mixed alcohol catalyst with 

subsequent separation. 

Stated that truck delivery of biomass was not 

practical for a 2,000 tonne/day facility. 

Train is now mentioned in Section 2.1 as an 

attractive alternative. 

Suggested adding the stoichiometric number to 

Table 8 and Table 10. 

Stoichiometric numbers have been added to 

tables 8 and 10. 

Suggested looking again at H2S concentrations in 

gas emitted from LO-CAT system. 

The H2S emissions were revisited and a 

reference is now provided. Additional polishing 

of effluent, if needed, is not expected to impact 

economics significantly. 

Suggested making continuous the biomass-togasoline process, without using a tank for the 

methanol. 

The design is set up for the process to be 

continuous with the possibility to store 

methanol prior to the MTG section. This is done 

in order to accommodate upsets or possibly 

buying methanol when market conditions allow 

for use during upsets to methanol production. 

104
 
 

 

 

 

 

 

Noted that LPG is marketable as is. Wording was changed to reflect this. 

Suggested adding an extra firewater pump and tank. Done. 

Suggested increasing the number of plant engineers 

from 1 to 2. 

Done. 

Suggested increasing the number of lab technicians 

from 2 to 4. 

Done. 

Somewhere in the report it would be good to Included throughout the report, and is 

highlight major opportunities to improve the cost of highlighted in the Executive Summary and 

BTG and where the major technology development Conclusions. 

challenges will lie. 

An nth plant means a mature industry and so 

feedstock is likely to be purpose-grown energy crop. 

Total cost of feedstock is expected to be higher than 

$50.7/ ton used based on our recent study. 

The feedstock cost is set in the OBP MYPP and 

remains $50.70/ton. However, a sensitivity is 

included for feedstock cost in Section 4.2. 

Even in a mature plant case, the project contingency 

as a percent of total installed cost is tiny compared to 

what most industrial companies use even for mature 

technologies. The assumed service factor of 96% 

seems high particularly for process that has 

significant amount of solids handling. 

There is a sensitivity included in Section 4.2 

addressing stream (service) factor. 

Recent peer/industry review of biomass gasification 

seems to indicate high costs. In particular I believe 

gas cleanup is an issue. Believe this review was 

published in December Energy issue. I haven’t seen 

this but NREL should check how that report 

compares to this and possibly acknowledge major 

differences if they exist. 

The cost of equipment, including gas cleanup 

equipment, is consistent with previous NREL 

reports. Efforts have been made at NREL to 

find appropriate equipment costs. Also, Section 

4.2, within the sensitivity analysis, shows effect 

on PGP if the total project investment is varied. 

Report uses thermal (BTU) basis to allocate prices to In Section 3, a market price for propane is given 

gasoline and LPG products. Alternate would be to and compared to the LPG price yielded by this 

use LPG market price and see what gasoline price method. 

results. 

Has MTG ever used syngas from wood gasification? In Section 5 it is stated that a future area of 

Carry over of contaminants or impurities are work would be to test the MTG reactor and 

potential problems for downstream catalysts. Is it catalyst with methanol from biomass-derived 

understood how to deal with this issue? Sensitivity syngas. 

to catalyst life should be tested. 

105
 
 

 

 

 

 

What is commercial status of fluid bed MTG? Issue 

is uncertainties around design, cost to build, and cost 

to operate. Fluid bed vs. fixed bed will have 

significant conversion differences that do not appear 

to be recognized. Fluid bed will not provide as much 

staging and will have lower yield than fixed bed. A 

sensitivity to using fixed vs. fluid bed should be 

included. 

A sensitivity for using fixed bed vs. fluid bed 

for a constant assumed conversion is included 

in Section 4.2 Methanol and Methanol-toGasoline. Compositions of several published 

MTG gasolines are compared with the Aspen 

simulation composition in Appendix B. 

The report cites a quote to build a facility to take 

natural gas to gasoline via methanol of $2.9B (2007 

$). Is there a reference for this? Assuming it is a 

valid quote, even after adjusting for plant size, this 

would suggest $1B for a 2,000 t/d size plant in the 

NREL report. Presumably the quote captures what 

was learned in the NZ demonstration plant. I 

wouldn’t expect an nth plant basis to reduce this cost 

by more than 50% ($0.5B) and it would take years 

of continuous improvement. I would also expect the 

investment cost to make syngas from wood to be 

higher than making syngas from natural gas. The 

NREL report shows an investment of $0.2 B which 

is significantly lower. This will be a major red flag 

for prospective participants in this technology. 

Reference of Seddon 2006 is given. The 

numbers were recalculated and the summary 

was rewritten to be more clear and concise. The 

New Zealand plant had several factors that 

contributed to higher costs in addition to being 

a first of a kind plant. See Appendix I. 

Figure 8 axis labeling does not match with 

individual figure in chart. The individual figures 

total 1.93 which corresponds to gasoline price per 

gallon. The axis refers to $/MM BTU. 

The axis of Figure 8 has been corrected. 

The sensitivity analysis starting on page 38 helps to The sensitivity parameters for TPI and 

highlight areas of uncertainty. NREL might say feedstock were increased. For feedstock the 

more about what the bases for variation are in the upper limit was $70/ton and it is now $85/ton. 

words that follow the table. Some of the ranges For TPI the upper limit was $230MM and now 

shown are low, especially feedstock and total project it is $500MM. 

investment. 

Some of this discussion should be included in the The Executive Summary now more explicitly 

executive summary. The summary should show explains that the PGP is for the base case and 

ranges of costs and yields...precision implied by the sensitivity analysis is provided for the 

single number with two significant figures reader to view the PGP changes with variations 

misrepresents the situation...is very misleading. in specified parameters. 

106
 
 

 

This process is very sensitive to the cost of 

feedstock...and therefore the overall yields. It would 

be useful to provide overview of yield assumptions 

for each part of plant and a comparison with actual 

currently demonstrated yields. 

Gasifier: Yields are well known based on 

experimental and published results. 

Tar Reformer: A sensitivity is presented in 

Sections 2.3 and 4.2 based on lower tar 

reformer conversions. 

Methanol Synthesis: Yields are well known and 

are based on published information. 

Methanol-to-Gasoline: A sensitivity is included 

in Section 4.2 on decreased gasoline yields. 

On p. 97 a comparison is made between PNNL’s and The comparison has been elaborated. See 

NREL’s report on BTG. I have not seen the PNNL Appendix I. 

report. The differences and consequences are 

indicated in the NREL report. In some cases the 

differences are pretty straightforward. However, the 

uninstalled cost difference is quite large, $84M for 

PNNL and $37M for NREL. More discussion about 

the bases and uncertainties which drive the 

differences and the implications for commercial 

readiness would be useful. 

107
 
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4. TITLE AND SUBTITLE

Gasoline from Wood via Integrated Gasification, Synthesis, and 

Methanol-to-Gasoline Technologies

5a. CONTRACT NUMBER

DE-AC36-08GO28308

5b. GRANT NUMBER

5c. PROGRAM ELEMENT NUMBER

6. AUTHOR(S)

S.D. Phillips, J.K. Tarud, M.J. Biddy, and A. Dutta 

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NREL/TP-5100-47594

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BB07.3710

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National Renewable Energy Laboratory

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14. ABSTRACT (Maximum 200 Words)

This report documents the National Renewable Energy Laboratory’s (NREL’s) assessment of the feasibility of making 

gasoline via the methanol-to-gasoline route using syngas from a 2,000 dry metric tonne/day (2,205 U.S. ton/day) 

biomass-fed facility. A new technoeconomic model was developed in Aspen Plus for this study, based on the model 

developed for NREL’s thermochemical ethanol design report (Phillips et al. 2007). The necessary process changes 

were incorporated into a biomass-to-gasoline model using a methanol synthesis operation followed by conversion, 

upgrading, and finishing to gasoline. Using a methodology similar to that used in previous NREL design reports and a 

feedstock cost of $50.70/dry ton ($55.89/dry metric tonne), the estimated plant gate price is $16.60/MMBtu 

($15.73/GJ) (U.S. $2007) for gasoline and liquefied petroleum gas (LPG) produced from biomass via gasification of 

wood, methanol synthesis, and the methanol-to-gasoline process. The corresponding unit prices for gasoline and 

LPG are $1.95/gallon ($0.52/liter) and $1.53/gallon ($0.40/liter) with yields of 55.1 and 9.3 gallons per U.S. ton of dry 

biomass (229.9 and 38.8 liters per metric tonne of dry biomass), respectively.

15. SUBJECT TERMS

biomass; gasoline; methanol; thermochemical; syngas; technoeconomic; advanced biofuels; gasification; biomass to 

gasoline; BTG; methanol to gasoline; MTG; plant gate price; ethanol.

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