Preliminary Screening — 

Technical and Economic 

Assessment of Synthesis Gas 

to Fuels and Chemicals with 

Emphasis on the Potential for 

Biomass-Derived Syngas 

December 2003 • NREL/TP-510-34929 

P.L. Spath and D.C. Dayton 

National Renewable Energy Laboratory 

1617 Cole Boulevard 

Golden, Colorado 80401-3393 

NREL is a U.S. Department of Energy Laboratory 

Operated by Midwest Research Institute • Battelle 

Contract No. DE-AC36-99-GO10337 
National Renewable Energy Laboratory 

1617 Cole Boulevard 

Golden, Colorado 80401-3393 

NREL is a U.S. Department of Energy Laboratory 

Operated by Midwest Research Institute • Battelle 

Contract No. DE-AC36-99-GO10337 

December 2003 • NREL/TP-510-34929 

Preliminary Screening — 

Technical and Economic 

Assessment of Synthesis Gas 

to Fuels and Chemicals with 

Emphasis on the Potential for 

Biomass-Derived Syngas 

P.L. Spath and D.C. Dayton 

Prepared under Task No. BBB3.4210 
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i

Executive Summary/Conclusions 

In principle, syngas (primarily consisting of CO and H2) can be produced from any 

hydrocarbon feedstock, including: natural gas, naphtha, residual oil, petroleum coke, 

coal, and biomass. The lowest cost routes for syngas production, however, are based on 

natural gas, the cheapest option being remote or stranded reserves. Economic 

considerations dictate that the current production of liquid fuels from syngas translates 

into the use of natural gas as the hydrocarbon source. Nevertheless, the syngas 

production operation in a gas-to-liquids plant amounts to greater than half of the capital 

cost of the plant. The choice of technology for syngas production also depends on the 

scale of the synthesis operation. Syngas production from solid fuels can require an even 

greater capital investment with the addition of feedstock handling and more complex 

syngas purification operations. The greatest impact on improving the economics of gasto liquids plants is through 1) decreasing capital costs associated with syngas production 

and 2) improving the thermal efficiency with better heat integration and utilization. 

Improved thermal efficiency can be obtained by combining the gas-to-liquids plant with a 

power generation plant to take advantage of the availability of low-pressure steam. 

The extensive research and development efforts devoted to syngas conversion to fuels 

and chemicals are documented in a vast amount of literature that tracks the scientific and 

technological advancements in syngas chemistry. The purpose of this report is to review 

the many syngas to products processes and summarize the salient points regarding the 

technology status and description, chemistry, catalysts, reactors, gas cleanliness 

requirements, process and environmental performances, and economics. Table 1 lists the 

products examined in this study and gives some facts about the technology as well as 

advantages and disadvantages. Table 2 summarizes the catalysts, process conditions, 

conversions, and selectivities for the various syngas to products processes. Table 3 

presents catalyst poisons for the various products. 

For all of the products examined in the economic analysis section (H2, MeOH, FTL, 

EtOH, mixed alcohols, olefins), syngas production accounts for at least 50% of the 

product cost and in many cases it is more like 75%. Overall, steps should be made to 

optimize the biomass-to-fuels process in order to obtain the highest yield, least cost 

configuration. To reduce costs, efforts should be focused on minimizing the cost of clean 

syngas production. The state of biomass gasifiers varies from the research and 

development stage to commercially available gasifiers. Research has been performed for 

various steps involved in cold and hot gas clean up and should be focused on 

demonstrating successful integrated low cost concepts. With the exception of mixed 

alcohols and ethanol, downstream syngas conversion technologies have all been 

demonstrated at the commercial scale. 

Overall, the information gathered in this report indicates that the best products to pursue 

are hydrogen and methanol. Ethanol from biomass-derived syngas could potentially be 

cost competitive. However, the state of this technology is still at the lab scale and the 

amount of data available is limited. Therefore, the ethanol analysis has a higher level of 

uncertainty than the other products examined. Additionally, because of the limited 
ii

amount of design data, more analysis should be performed for mixed alcohols synthesis 

to examine biomass-optimized configurations including recycle for maximum conversion 

and the resulting economics. 

All biomass fuels have potential to significantly reduce the import of petroleum products. 

Additionally, economies of scale can play a large factor in lowering the product cost. 

Therefore, opportunities to co-feed with coal or natural gas systems may be one way to 

get renewable fuels into the marketplace, just as co-firing biomass with coal is being 

done in the power generation industry. 

The impetus for this extensive literature review is the recent re-organization of the U.S. 

Department of Energy’s Office of Energy Efficiency and Renewable Energy (EERE) into 

eleven new Program Offices. EERE’s mission is to strengthen US energy security, 

environmental quality, and economic health. The Office of Biomass Program is one of 

these newly created offices. Its goal is developing technologies to transform our 

abundant biomass resources into clean, affordable, and domestically-produced biofuels, 

biopower, and high-value bioproducts resulting in economic development, energy supply 

options, and energy security. In this context, commercially available and nearcommercial syngas conversion processes were evaluated on technological, 

environmental, and economic bases. This report serves as a first step. Additional, more 

detailed analyses are required to identify promising, cost-effective fuel synthesis 

technologies where biomass thermochemical conversion could make an impact. 
iii

Table 1: Summary of Product Information – Facts, Advantages, & Disadvantages 

Product Facts Advantages Disadvantages H2  Largest use of syngas 

 Predominately made via SMR 

 Appears to be the most cost competitive 

option for biomass 

 Automakers working on H2 fueled 

vehicles 

 H2 itself is a clean fuel 

 Concentrated CO2 stream 

 Mature technology  High yields 

 Delivered price can b

than plant gate pric

FTL  Existing plants in South Africa, 

Malaysia, and New Zealand 

 Predominately made from coal but 

South Africa switching some operating plants to natural gas 

 Many companies are examining FTL 

for stranded natural gas 

 Yields petroleum products consistent with 

today’s fuel supply 

 Successful operation of slurry phase 

reactor 

 No sulfur and low aromatic products  Reduced tailpipe emissions compared to 

conventional diesel 

 Low yields  Unavoidable broad ra

selectivity 

 Need to use all strea

making it complex (more than 200 fuel 

 Currently, only econo

petroleum is not pre

NH3  Largest consumer of H2

 2nd largest synthetic chemical product 

 Mature technology  Toxic 

MeOH  Predominately from syngas made via 

SMR 

 Several automakers have developed 

prototype direct MeOH fuel cell vehicles 

 Commodity chemical 

 Mature technology  Many uses for MeOH (formaldehyde, 

acetic acid, DME, MTBE, MTG, MTO, MOGD) 

 Fuel used as M100 and M85  Possible source of H2 & MeOH reforming 

is a practiced technology 

 Have to compete with  LPMeOH still in deve Poor solubility in gaso

separation problem

 More corrosive than g M85 has not develop

DME  Many potential uses (diesel or 

cooking fuel, refrigerant, chemical feedstock) 

 In situ dehydration of MeOH to DME 

increases MeOH yields 

Acetic acid  Half of worldwide production comes 

from MeOH carbonylation 

Formaldehyde  Largest consumer of MeOH 

 Demand driven by construction 

industry 

MTBE  95% used in gasoline pool 

 North America consumes 65% of 

worldwide production 

 Environmental conce

contamination 
iv

Product Facts Advantages Disadvantages MTO and MOGD  Low aromatic content in MOGD gasoline 

 Selectivity of olefins to gasoline & diesel > 

95% for MOGD 

 No commercial MTO 

UOP and HYDRO dthat produces ethyl

MTG  One plant in New Zealand built in 

1985 but currently it produces only MeOH 

 High quality, high octane gasoline  High amount of arom

 High decalene conce

TIGAS  MeOH to gasoline  No commercial plants 

 TIGAS yields a lower 

MTG 

Ethanol  Currently produced primarily via direct 

fermentation of carbohydrates 

 Is currently a fuel additive  Currently, has a tax credit  Compared to direct fermentation, able to 

process nearly any biomass resource 

 No commercial plants  Many universities hav

designs but presenis unknown 

Mixed alcohols  Alcohol fuels developed by many 

 Some tested in Europe 

 More attractive blending stock than 

MeOH 

 Lower H2:CO ratio required 

 No stand alone comm Low yields  Poor selectivity  Currently, there is litt

alcohol synthesis 

Oxosynthesis products 

 Hydroformylation of olefins with 

syngas 

 Produces C3-C15 aldehydes which are 

typically converted to alcohols and acids 

 Lower H2:CO ratio required 

 Very versatile process  Commercial process 

 Product/catalyst sepa Limited to high value 

Isosynthesis products 

 Converts syngas to isobutene and 

isobutane 

 Lower H2:CO ratio required  No commercial plants 

 Low catalyst activity  Must compete with is
v

Table 2: Summary of Syngas Conversion Processes and Conditions 

Process Catalyst Process Conditions 

T (°C) P (bar) H2:CO 

% conv (CO basis) 

Product Se

Fe 300-350 10-40 1.7:1 α-olefins 

gasoline 

ASF 15-

Co 200-240 7-12 2.15:1 

50-90% with recycle 

Waxes diesel 

ASF 

Fischer-Tropsch Synthesis 

Ru Waxes 

ZnO/Cr2O3 350 250-350 3:1 

Methanol Synthesis Cu/ZnO/Al2O3 220-275 50-100 

99% (25% max/pass – 4-7% actual/pass) 

Methanol 

> 9

Ammonia Fe/FeO + additives 430-480 

(550 max) 

100-500 2-3:1 H2:N2 10-35%/pass Ammonia > 9

Alkali/ZnO/Cr2O3 300-425 125-300 1:1 5-20% Branched 

primary alcohols 

Alkali/Cu/ZnO(Al2O3) 275-310 50-100 2-3:1 20-30% Primary 

alcohols 

30-

17-

Alkali/CuO/CoO 260-340 60-200 0.5-4:1 5-30% Linear primary 

alcohols 

ASF 

Higher Alcohol Synthesis 

Alkali/MoS2 260-350 30-175 1:1 10% Linear 

alcohols 

75-

pro

Co carbonyl 110-200 200-300 1:1 + olefin Co – P modified 160-200 50-100 1:1 + olefin C11-C14

Oxosynthesis alcohols 

Rh – P modified 60-120 7-25 1:1 + 

propylene 

 C4 aldehydes > 9

Isosynthesis ThO2 400-450 100-1000 

(300) 

0.85:1 40-50% 

 ZrO2 300-425 350 1:1 30% 

i-C4

15 

Steam Methane Reforming 

Ni 850 15-30 na 100% CH4

conversion 

Syngas/ hydrogen 
vi

Table 3: Summary of Poisons 

Process Contaminant Level Source/comments 

Sulfur 0.2 ppm 

1 ppmv 60 ppb 

Dry, 1981 Boerrigter, et al, 2002 Turk, et al, 2001 

Fischer-Tropsch Synthesis Halides 10 ppb Boerrigter, et al, 2002 

Nitrogen 10 ppmv NH3 

0.2 ppmv NOx 10 ppb HCN 

Turk, et al, 2001 

Sulfur (not COS) 

<0.5 ppmv (<0.1 ppmv preferred) 

Kung, 1992 

Halides 0.001 ppmv Twigg and Spencer 2001 Methanol Synthesis 

Fe and Ni 0.005 ppmv Kung, 1992 

Sulfur (including OCS) 0.1 ppmv Total halides 0.01 ppmv Acetylene 5 ppmv Total unsaturates 300 ppmv NH3 10 ppmv HCN 0.01 ppmv 

LPMeOH Synthesis 

Fe and Ni 0.01 ppmv 

Novem (2002) 

H2O 200 ppm Revsible CO 200 ppm Revsible CO2 100 ppm Revsible O2 100 ppm Revsible Sulfur (H2S) 0.1 ppm Irrevsible Chlorine 0.1 ppm Irrevsible 

Ammonia Synthesis 

As, P, Sb --- Irrevsible 

Ethanol Synthesis Very little work has been published on the effects of syngas impurities Higher Alcohol Synthesis - modified FT catalysts are the same as those for FT catalyst - modified methanol synthesis catalysts are the same as those for meOH catatysts Oxosynthesis Strong acids, HCN, organosulfur, H2S, COS, O2, and dienes (Bahrmann and Bach 2000) Isosynthesis Thoria catalysts are not poisoned by sulfur and have high resistance to other poisons as weSteam Methane Reforming <0.5 ppm for reformer catalyst life of 3 years 
vii

Acknowledgements 

The authors acknowledge the following colleagues from the National Renewable Energy 

Laboratory who contributed information and assisted in the economic assessment: Matthew 

B. Ringer, John L. Jechura, and Margaret K. Mann. Additionally, we acknowledge Tim 

Eggeman from Neoterics International who contributed the ethanol section of this report. 
viii

Table of Contents 

Executive Summary/Conclusions ........................................................................................... i

Acknowledgements ............................................................................................................... vii

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

2 Hydrogen .............................................................................................................................. 5

2.1 Summary.......................................................................................................................... 5 

2.2 Introduction ..................................................................................................................... 7 

2.3 Technology Description .................................................................................................. 7 

2.4 Hydrogen From Other Compounds ................................................................................. 8 

2.5 Chemistry ........................................................................................................................ 8 

2.6 Catalysts .......................................................................................................................... 9 

2.7 Reactors ......................................................................................................................... 10 

2.8 Gas Cleanliness Requirements ...................................................................................... 10 

2.9 Status/Technology Developers...................................................................................... 11 

2.10 Process Performance ................................................................................................... 12 

2.11 Environmental Performance........................................................................................ 12 

2.12 Cost.............................................................................................................................. 13 

2.13 R&D Needs ................................................................................................................. 14 

2.14 References ................................................................................................................... 14 

3 Ammonia............................................................................................................................. 16

3.1 Summary........................................................................................................................ 16 

3.2 Introduction ................................................................................................................... 18 

3.3 Technology Description ................................................................................................ 19 

3.4 Chemistry ...................................................................................................................... 20 

3.5 Catalysts ........................................................................................................................ 21 

3.6 Reactors ......................................................................................................................... 23 

3.7 Gas Cleanliness Requirements ...................................................................................... 25 

3.8 Status/Technology Developers...................................................................................... 26 

3.9 Environmental Performance.......................................................................................... 27 

3.10 Cost.............................................................................................................................. 27 

3.11 R&D Needs ................................................................................................................. 28 

3.12 References ................................................................................................................... 28 
ix

4 Methanol and Methanol Derivative (including formaldehyde, acetic acid, MTG, 

MTO, MOGD, TIGAS, and DME)................................................................................... 29

4.1 Summary........................................................................................................................ 29 

4.2 Introduction ................................................................................................................... 31 

4.3 Technology Description ................................................................................................ 31 

4.4 Reactors ......................................................................................................................... 32 

4.5 Chemistry ...................................................................................................................... 34 

4.6 Catalysts ........................................................................................................................ 35 

4.7 Gas Cleanliness Requirements ...................................................................................... 36 

4.8 Status/Technology Developers...................................................................................... 38 

4.9 Process Performance ..................................................................................................... 39 

4.10 Uses ............................................................................................................................. 40 

4.11 Environmental Performance........................................................................................ 47 

4.12 Cost of Methanol ......................................................................................................... 47 

4.13 R&D Needs ................................................................................................................. 51 

4.14 References ................................................................................................................... 52 

5 Ethanol................................................................................................................................ 58

5.1 Summary........................................................................................................................ 58 

5.2 Introduction ................................................................................................................... 60 

5.3 Technology Description ................................................................................................ 61 

5.4 Chemistry And Microbiology........................................................................................ 63 

5.5 Fermentor Design .......................................................................................................... 65 

5.6 Gas Cleanliness Requirements ...................................................................................... 65 

5.7 Status/Technology Developers...................................................................................... 65 

5.8 Environmental Performance.......................................................................................... 66 

5.9 Cost................................................................................................................................ 67 

5.10 R&D Needs ................................................................................................................. 68 

5.11 References ................................................................................................................... 69 

6 Mixed Higher Alcohols...................................................................................................... 70

6.1 Summary........................................................................................................................ 70 

6.2 Introduction ................................................................................................................... 72 

6.3 Technology Description ................................................................................................ 73 

6.4 Chemistry ...................................................................................................................... 73 
x

6.5 Catalysts ........................................................................................................................ 76 

6.6 Reactors ......................................................................................................................... 79 

6.7 Gas Cleanliness Requirements ...................................................................................... 79 

6.8 Status/Technology Developers...................................................................................... 79 

6.9 Environmental Performance.......................................................................................... 83 

6.10 Cost.............................................................................................................................. 83 

6.11 R&D Needs ................................................................................................................. 83 

6.12 References ................................................................................................................... 85 

7 Fischer-Tropsch Synthesis ................................................................................................ 90

7.1 Summary........................................................................................................................ 90 

7.2 Introduction ................................................................................................................... 92 

7.3 Technology Description ................................................................................................ 93 

7.4 Chemistry ...................................................................................................................... 94 

7.5 Catalysts ........................................................................................................................ 96 

7.6 Reactors ......................................................................................................................... 97 

7.7 Gas Cleanliness Requirements ...................................................................................... 99 

7.8 Status/Technology Developers.................................................................................... 101 

7.9 Process Performance ................................................................................................... 102 

7.10 Environmental Performance...................................................................................... 102 

7.11 Cost............................................................................................................................ 103 

7.12 R&D Needs ............................................................................................................... 106 

7.13 References ................................................................................................................. 107 

8 Oxosynthesis ..................................................................................................................... 111

8.1 Summary...................................................................................................................... 111 

8.2 Introduction ................................................................................................................. 113 

8.3 Technology Description .............................................................................................. 113 

8.4 Chemistry .................................................................................................................... 114 

8.5 Catalysts ...................................................................................................................... 115 

8.6 Gas Cleanliness Requirements .................................................................................... 116 

8.7 Status/Technology Developers.................................................................................... 116 

8.8 Cost.............................................................................................................................. 119 

8.9 R&D Needs ................................................................................................................. 119 

8.10 References ................................................................................................................. 120 
xi

9 Isosynthesis ....................................................................................................................... 121

9.1 Summary...................................................................................................................... 121 

9.2 Introduction ................................................................................................................. 123 

9.3 Technology Description .............................................................................................. 123 

9.4 Chemistry .................................................................................................................... 123 

9.5 Catalysts ...................................................................................................................... 124 

9.6 Gas Cleanliness Requirements .................................................................................... 125 

9.7 Status/Technology Developers.................................................................................... 125 

9.8 R&D Needs ................................................................................................................. 125 

9.9 References ................................................................................................................... 125 

10 Yields, Energy, and Economics..................................................................................... 127

10.1 Summary.................................................................................................................... 127 

10.2 Product Yields ........................................................................................................... 129 

10.3 Life Cycle Energy Ratios .......................................................................................... 131 

10.4 Petroleum Product Replacement ............................................................................... 133 

10.5 Economics ................................................................................................................. 135 

10.6 References ................................................................................................................. 141 
xii

List of Figures 

Figure 1: Diagram of Syngas Conversion Processes Covered in This Report ........................ 3 

Figure 2: Block Flow Diagram of Hydrogen via Steam Methane Reforming......................... 8 

Figure 3: Generic Ammonia Synthesis Process Flow Diagram............................................. 19 

Figure 4: Various Ammonia Converter Designs ................................................................... 24 

Figure 5: Simplified Methanol Synthesis Process Flow Diagram......................................... 31 

Figure 6: Syngas to Ethanol Simplified Block Flow Diagram .............................................. 61 

Figure 7: FTS General Process Flow Diagram...................................................................... 93 

Figure 8: Anderson-Shulz-Flory Distribution........................................................................ 95 

Figure 9: Types of Fischer-Tropsch Synthesis Reactors ....................................................... 98 

Figure 10: Schematic Flow Diagram of the LP Oxo Process .............................................. 117 

Figure 11: Product Yields from Fossil and Biomass Resources.......................................... 130 

Figure 12: Life Cycle Fossil Energy Ratio .......................................................................... 132 

Figure 13: Life Cycle Primary Energy Ratio....................................................................... 132 

Figure 14: Potential Replacement of Petroleum Products with Biomass Fuels 

(Energy Basis)..................................................................................................... 134 

Figure 15: Product Price in $/GJ (LHV).............................................................................. 139 

Figure 16: Capital Cost Breakdown for Biomass to MeOH Systems.................................. 140 
xiii

List of Tables 

Table 1: Summary of Product Information – Facts, Advantages, & 

Disadvantages ...........................................................................................................iii 

Table 2: Summary of Syngas Conversion Processes and Conditions...................................... v 

Table 3: Summary of Poisons................................................................................................. vi

Table 4: Catalytic Steam Reforming Licensors*................................................................... 11 

Table 5: Commercial Suppliers of PSA Units** ................................................................... 12 

Table 6: Comparison of Biomass Based Hydrogen Studies .................................................. 13 

Table 7: Hydrogen Cost Data From other Sources................................................................ 14 

Table 8: Summary of Commercial Ammonia Synthesis Catalyst Compositions.................. 23 

Table 9: Summary of Ammonia Synthesis Catalyst Poisons ................................................ 26 

Table 10: Comparison of Ammonia Economics.................................................................... 27 

Table 11: Commercial Methanol Synthesis Catalyst Formulations ...................................... 36 

Table 12: MeOH Gaseous Contaminant Constraints............................................................. 38 

Table 13: Methanol Technology Suppliers............................................................................ 39 

Table 14: Conversions for Methanol, Methanol/DME, and DME ........................................ 45 

Table 15: Comparison of Coal Based Methanol Economic Studies...................................... 48 

Table 16: Comparison of Biomass Based Methanol Studies................................................. 49 

Table 17: Comparison of Formaldehyde Production Costs................................................... 50 

Table 18: Comparison of Coal Based MTG Economic Studies ............................................ 51 

Table 19: Ethanol Yields on a Common Basis ...................................................................... 63 

Table 20: Summary of USDA Survey Result (Cash Costs in $/gal of ethanol) ..................... 68 

Table 21: Companies and Information Regarding Mixed Alcohols Research ...................... 82 

Table 22: EcaleneTM Composition......................................................................................... 83 

Table 23: Comparison of Mixed Higher Alcohol Economics From Natural Gas ................. 84 

Table 24: Syngas Impurities and Tolerances for Fischer-Tropsch ...................................... 100 

Table 25: Price Ranges for Natural Gas, Gasoline, and Diesel ........................................... 104 

Table 26: Comparison of Biomass Based Fischer Tropsch Studies .................................... 106 

Table 27: Summary of U.S. Oxo Chemical Producers ........................................................ 118 

Table 28: U.S. Prices for Oxo Chemicals............................................................................ 119 

Table 29: Feedstock to Product Material Balance Information........................................... 129 

Table 30: Life Cycle Energy Ratios (LHV basis)................................................................ 131 
xiv

Table 31: Economic Parameters for Discounted Cash Flow Analysis ................................ 136 

Table 32: Summary of Cases Examined.............................................................................. 137 

Table 33: Base Case Minimum Product Selling Price......................................................... 138 

Table 34: Price of Commercial Products............................................................................. 138 

Table 35: Sensitivity – Minimum Product Selling Price ..................................................... 139 
1

1 Introduction 

In its simplest form, syngas is composed of two diatomic molecules, CO and H2 that 

provide the building blocks upon which an entire field of fuel science and technology is 

based. Over the years, the gaseous mixture of CO and H2 has had many names depending 

on how it was formed; producer gas, town gas, blue water gas, synthesis gas, and syngas, 

to name a few. In the 1800s coal gasification was used to provide much of the syngas 

used for lighting and heating. The beginning of the 20th century saw the dawn of fuels 

and chemicals synthesis from syngas. 

The synthesis of hydrocarbons from CO hydrogenation was discovered in 1902 by 

Sabatier and Sanderens who produced methane by passing CO and H2 over Ni, Fe, and 

Co catalysts. At about the same time, the first commercial hydrogen from syngas 

produced from steam methane reforming was commercialized. Haber and Bosch 

discovered the synthesis of ammonia from H2 and N2 in 1910 and the first industrial 

ammonia synthesis plant was commissioned in 1913. The production of liquid 

hydrocarbons and oxygenates from syngas conversion over iron catalysts was discovered 

in 1923 by Fischer and Tropsch. Variations on this synthesis pathway were soon to 

follow for the selective production of methanol, mixed alcohols, and isosynthesis 

products. Another outgrowth of Fischer-Tropsch Synthesis (FTS) was the 

hydroformylation of olefins discovered in 1938. 

Many of the syngas conversion processes were developed in Germany during the first 

and second world wars at a time when natural resources were becoming scacre and 

alternative routes for hydrogen production, ammonia synthesis, and transportation fuels 

were a necessity. Thus the poor economics of some of these syngas conversion processes 

were of little consequence. With the development of the petroleum industry in the 1940s 

and beyond, the unattractive economics of many of these syngas routes became an issue 

and were replaced by petroleum-based processes. Methanol and ammonia continue to be 

produced from syngas using similar processes. Apart from hydrogen production, these 

processes constitute the major uses of syngas. 

While petroleum-derived transportation fuels and chemicals were taking center stage in 

the worldwide market, certain political and environmental drivers were helping to 

develop improved commercial syngas conversion processes. The political situation in 

South Africa and the abundance of local coal reserves in that country helped bring about 

the most successful commercial FTS industry in the world based on syngas production 

from coal gasification. Sasol currently supplies diesel, gasoline, and other high-value 

hydrocarbons to local and global markets. The Arab oil embargo in the 1970s also 

highlighted the US dependence on foreign oil imports and helped renew interest in 

syngas conversion technologies. 

Increasing environmental concerns and tighter regulations surrounding fossil fuel use also 

provided impetus for syngas conversion technologies to produce cleaner (virtually no 

sulfur) fuels and chemicals. The use of methanol and isobutene for the production of 

methyl tert-butyl ether (MTBE), an octane enhancing oxygenated component in 

reformulated gasoline, also increased demand for syngas conversion technologies. 
2

MTBE itself, however, is becoming an environmental concern as a watershed pollutant 

and the future use of this oxygenate is uncertain. The latest environmental driver to 

likely increase demand for syngas even more is the goal of establishing a hydrogen 

economy. The vision is that hydrogen will be the fuel of choice for transportation and 

electricity generation via high efficiency, environmentally benign fuel cells. 

In principle, syngas can be produced from any hydrocarbon feedstock. These include: 

natural gas, naphtha, residual oil, petroleum coke, coal, and biomass. The lowest cost 

routes for syngas production, however, are based on natural gas. The cheapest option is 

remote or stranded reserves. Current economic considerations dictate that the production 

of liquid fuels from syngas translates into using natural gas as the hydrocarbon source. 

Nevertheless, the syngas production operation in a gas-to-liquids plant amounts to greater 

than half of the capital cost of the plant. The choice of technology for syngas production 

also depends on the scale of the synthesis operation. Syngas production from solid fuels 

can require an even greater capital investment with the addition of feedstock handling 

and more complex syngas purification operations. The greatest impact on improving gasto-liquids plant economics is to decrease capital costs associated with syngas production 

and improve thermal efficiency through better heat integration and utilization. Improved 

thermal efficiency can be obtained by combining the gas-to-liquids plant with a power 

generation plant to take advantage of the availability of low-pressure steam. 

The syngas composition, most importantly the H2/CO ratio, varies as a function of 

production technology and feedstock. Steam methane reforming yields H2/CO ratios of 

3/1 while coal gasification yields ratios closer to unity or lower. Conversely, the 

required properties of the syngas are a function of the synthesis process. Fewer moles of 

product almost always occur when H2 and CO are converted to fuels and chemicals. 

Consequently, syngas conversion processes are more thermodynamically favorable at 

higher H2 and CO partial pressures. The optimum pressures depend on the specific 

synthesis process. 

With the exception of methane steam reforming, catalytic syngas conversion processes 

are exothermic reactions generating large excesses of heat. This highlights the specific 

need for removing this heat of reaction to carefully control reaction temperatures and 

maintain optimized process conditions. Maximizing product yields, minimizing side or 

competing reactions, and maintaining catalyst integrity dictate optimum synthesis 

reaction temperatures. 

Catalysts play a pivotal role in syngas conversion reactions. In fact, fuels and chemicals 

synthesis from syngas does not occur in the absence of appropriate catalysts. The basic 

concept of a catalytic reaction is that reactants adsorb onto the catalyst surface and 

rearrange and combine into products that desorb from the surface. One of the 

fundamental functional differences between syngas synthesis catalysts is whether or not 

the adsorbed CO molecule dissociates on the catalyst surface. For FTS and higher 

alcohol synthesis, CO dissociation is a necessary reaction condition. For methanol 

synthesis, the CO bond remains intact. Hydrogen has two roles in catalytic syngas 

synthesis reactions. In addition to being a reactant needed for CO hydrogenation, it is 

commonly used to reduce the metalized synthesis catalysts and activate the metal surface. 
3

Since the genesis of syngas conversion to fuels and chemicals, a tremendous amount of 

research and development has been devoted to optimizing product yields and process 

efficiencies. This includes the discovery of catalysts with optimized formulations 

containing the most active metals in combination with appropriate additives to improve 

activity and selectivity in a given process. Mechanistic studies have been conducted to 

interpret the fundamentals of specific conversion processes and measure the kinetic rates 

of key chemical reactions. Reactor design and engineering is another active research and 

development area of syngas conversion technology. Temperature control and stability in 

conversion reactors is a critical process parameter because of the large excess heat of 

reaction. Detailed process engineering and integration with respect to heat integration 

and syngas recycle to improve conversion efficiencies is used to optimize commercial 

synthesis processes. 

Given the rich history of syngas conversion and the extensive research and development 

efforts devoted to this field of study, it is not surprising that a vast amount of literature is 

available that tracks the scientific and technological advancements in syngas chemistry. 

The purpose of this report is to review the many syngas to products processes and 

summarize the salient points regarding the technology status and description, chemistry, 

catalysts, reactors, gas cleanliness requirements, process and environmental performances, and 

economics. A diagram highlighting the specific processes reviewed in this report is 

shown in Figure 1. 

Figure 1: Diagram of Syngas Conversion Processes Covered in This Report 

Syngas

CO + H2

Methanol

H2O

WGS

Purify

H2

N2 over Fe/FeO

(K2O, Al2O3, CaO)

NH3

Cu/ZnO Isosynthesis

ThO2 or ZrO2

i-C4

Alkali-doped

ZnO/Cr2O3

Cu/ZnO; Cu/ZnO/Al2O3

CuO/CoO/Al2O3

MoS2

Mixed

Alcohols

Oxosynthesis

HCo(CO)4

HCo(CO)3P(Bu3)

Rh(CO)(PPh3)3 Aldehydes

Alcohols

Fischer-Tropsch

Fe, C

o, Ru

Waxes

Diesel

Olefins

Gasoline

Ethanol

Co, Rh

Formaldehyde

Ag

DMEAl

O2 3

zeolites

MTO

MTG

Olefins

Gasoline

MTBE

Acetic Acid

carbonylation

CH3OH + CO

Co, Rh, Ni

M100

M85

DMFC

Direct Use

homologation

Co

isobutyleneacidic ion exchange
4

The impetus for this extensive literature review is the recent re-organization of the U.S. 

Department of Energy’s Office of Energy Efficiency and Renewable Energy (EERE) into 

eleven new Program Offices. EERE’s mission is to strengthen US energy security, 

environmental quality, and economic health. The Office of Biomass Program is one of 

these newly created offices. Its goal is developing technologies to transform our 

abundant biomass resources into clean, affordable, and domestically produced biofuels, 

biopower, and high-value bioproducts resulting in economic development, energy supply 

options, and energy security. In this context, commercially available and nearcommercial syngas conversion processes were evaluated on technological, 

environmental, and economic bases. This report serves as a first step. Additional, more 

detailed analyses are required to identify promising, cost-effective fuel synthesis 

technologies where biomass thermochemical conversion could make an impact. 
5

2 Hydrogen 

2.1 Summary 

Hydrogen production is the largest use of syngas. Hydrogen is mainly consumed for 

ammonia production, followed by refining and methanol production. Hydrogen is 

primarily produced from methane. 

The dominant technology for hydrogen production is steam methane reforming. The 

process can be divided into the following four steps: feed pretreatment, steam reforming, 

CO shift conversion, and hydrogen purification. Reforming (CnHm + nH2O Æ

(n+m/2)H2 + nCO) and shift (CO + H2O Æ CO2 + H2) are the main reactions. If the 

feedstock is methane then 50% of the hydrogen comes from the steam. The reforming 

reaction is highly endothermic and is favored by high temperatures and low pressures. 

The shift reaction is exothermic and favors low temperatures. In industrial reformers, the 

reforming and shift reactions result in a product composition that closely approaches 

equilibrium. Presently, most plants use pressure swing adsorption (PSA) after the shift 

conversion step, to purify the hydrogen. Using PSA, a hydrogen purity greater than 

99.99% can be achieved. 

BASF was the first company to develop steam methane reforming in the early 1900s 

(Nirula, 1995). Today there are many licensors for the various components of the steam 

reforming process. A wide variety of reformer designs exist and can be used in various 

process configurations. The reformer steam to carbon ratio is usually between 2-6 

depending on the feedstock and process conditions. Excess steam is used to prevent 

coking in the reformer tubes. Conventional steam reforming catalysts are 10-33 wt% 

NiO on a mineral support (alumina, cement, or magnesia). The high temperature shift 

(HTS) catalyst has an iron oxide, chromium oxide basis while the major component in 

the low temperature shift (LTS) catalyst is copper oxide, usually in a mixture with zinc 

oxide. 

Sulfur compounds are the main poison of reforming catalyst. Even at a concentration of 

0.1 ppm the catalyst can begin to deactivate. The LTS catalyst is also very sensitive to 

sulfur. Additionally, chloride is a LTS catalyst poison. HTS catalysts can tolerate sulfur 

concentrations up to several hundred ppm, although the activity will decline. Other HTS 

catalyst poisons include phosphorus, silicon and unsaturated hydrocarbons in the 

presence of NOx (Häussinger, 2000). Prior to the PSA unit, entrained liquids (water and 

condensed hydrocarbons) must be removed because they will permanently damage the 

adsorbents. 

Hydrogen itself is a clean burning fuel. However, depending upon the feedstock used, its 

production can generate a considerable amount of CO2. Additionally, steam reformers 

produce NOx from fuel combustion. Generally, as the feedstock goes from natural gas to 

light hydrocarbons to heavy hydrocarbons and then to solid feedstocks, there is an 

increase in processing difficulty and capital costs. Merchant hydrogen prices vary 

considerably depending on the volume and form of delivery. For large volume users, 

pipeline delivery is the most economical followed by bulk liquid hydrogen delivery. 
6

Section Contents 

2.1 Summary .......................................................................................................................... 5

2.2 Introduction ..................................................................................................................... 7

2.3 Technology Description................................................................................................... 7

2.4 Hydrogen From Other Compounds................................................................................. 8

2.5 Chemistry......................................................................................................................... 8

2.6 Catalysts.......................................................................................................................... 9

2.7 Reactors......................................................................................................................... 10

2.8 Gas Cleanliness Requirements...................................................................................... 10

2.9 Status/Technology Developers ...................................................................................... 11

2.10 Process Performance................................................................................................... 12

2.11 Environmental Performance........................................................................................ 12

2.12 Cost.............................................................................................................................. 13

2.13 R&D Needs.................................................................................................................. 14

2.14 References.................................................................................................................... 14

List of Figures 

Figure 2: Block Flow Diagram of Hydrogen via Steam Methane Reforming......................... 8 

List of Tables 

Table 4: Catalytic Steam Reforming Licensors* ...................................................................... 11 

Table 5: Commercial Suppliers of PSA Units**.................................................................... 12 

Table 6: Comparison of Biomass Based Hydrogen Studies.................................................. 13 

Table 7: Hydrogen Cost Data From other Sources............................................................... 14 
7

2.2 Introduction 

The largest use of syngas is for hydrogen production (Wender, 1996). Hydrogen is 

produced as both a main product and as a by-product. Hydrogen producers often 

consume the product captively (e.g., ammonia producers, oil refineries, and methanol 

producers). The total amount of hydrogen consumed worldwide in 1999 was 15,864 

billion ft3; 20% of this was consumed in the U.S. (Suresh, et al, 2001). The majority of 

the worldwide hydrogen consumption (about 60%) was for ammonia production. This 

was followed by refining at 23% and methanol production at 9%, leaving only 8% as 

merchant hydrogen. 

Currently, hydrogen is primarily produced from methane. However, other hydrocarbon 

feedstocks include naphtha, heavy residues from petrochemical processes, coke oven gas, 

and coal. Electrolysis has also been used to produce hydrogen primarily where there is 

cheap electricity from hydropower stations. Presently, 77% of the worldwide hydrogen 

production comes from petrochemicals, 18 % from coal, 4 % from water electrolysis, and 

1% from other sources (Häussinger, et al, 2000). The flexibility of hydrogen to be 

produced from any resource enables production from locally available, lowest-cost 

resources. 

2.3 Technology Description 

The dominant technology for hydrogen production is currently steam methane reforming. 

The process can be divided into the following four steps: feed pretreatment, steam 

reforming, CO shift conversion, and hydrogen purification. For natural gas, the only 

pretreatment required is desulfurization, which usually consists of a hydrogenator 

followed by a zinc oxide bed. After desulfurization, natural gas is fed to a reformer 

reactor, where it reacts with steam to produce CO, CO2, and H2 through the reforming 

(CH4 + H2 Æ CO + 3H2) and water gas shift (CO + H2O Æ CO2 + H2) reactions. The 

reformer reactor is comprised of catalyst-filled tubes, surrounded by a firebox that 

provides the heat necessary for the endothermic reforming reaction which operates at 

about 850°C and between 1.5-3 MPa. The gas exiting the reformer is cooled to about 

350°C then undergoes the water gas shift (WGS) reaction in a high temperature shift 

(HTS) converter. 

In the “classical process”, the gas is further cooled to about 220°C and undergoes WGS 

in a low temperature shift (LTS) converter followed by CO2 scrubbing (e.g., 

monoethanolamine or hot potash). Finally, to remove trace amounts of CO and CO2 a 

methanation reactor is used (CO + 3H2 Æ CH4 + H2O). The product hydrogen has a 

purity of 97-99%. 

Currently, after the shift conversion step, hydrogen is purified using a pressure swing 

adsorption (PSA) unit to obtain purity greater than 99.99%. Since the PSA unit can 

easily remove CO and other components to produce a high purity hydrogen stream, most 

of today’s hydrogen plants use only the HTS reactor (Leiby, 1994). However, the LTS 

reactor can increase the hydrogen yield slightly. The PSA off-gas, which contains 
8

unreacted CH4, CO, and CO2, plus unrecovered hydrogen, is used to fuel the reformer. 

This stream usually supplies 80-90% of the reformer heat duty supplemented by natural 

gas to fulfill the remaining heat requirement (Leiby, 1994). Figure 2 depicts this process. 

Figure 2: Block Flow Diagram of Hydrogen via Steam Methane Reforming 

2.4 Hydrogen From Other Compounds 

In addition to the direct production of hydrogen from gaseous hydrocarbon feedstocks, 

hydrogen can be produced from liquid energy carriers such as ethanol and methanol as 

well as from ammonia. Methanol reforming (CH3OH + H2O Æ CO2 + 3H2) is practiced 

in Japan and, to a lesser degree, in Europe where there are no economical sources of 

syngas (Suresh, et al, 2001). Currently, there are about a dozen companies that are 

involved in building plants to produce hydrogen from methanol or ammonia. 

2.5 Chemistry 

Steam reforming of hydrocarbons involves the catalytic conversion of a hydrocarbon 

feedstock and steam to hydrogen and carbon oxides. Generally speaking, the chemical 

process for steam reforming of hydrocarbons is described by the following equation 

(Häussinger, et al, 2000): 

CnHm + nH2O Æ (n+m/2)H2 + nCO 

Water gas shift is another important reaction that occurs in the reformer. For steam 

methane reforming the following two reactions occur in the reformer (Leiby, 1994): 

CH4 + H2O Æ 3H2 + CO ∆Hr = 49.3 kcal/mol 

CO + H2O Æ H2 + CO2 ∆Hr = -9.8 kcal/mol 

Therefore, in this case, 50% of the hydrogen comes from the steam. The reforming 

reaction is highly endothermic and is favored by high temperatures and low pressures. 
9

Higher pressures tend to lower the methane conversion. In industrial reformers the 

reforming and shift reactions result in a product composition that closely approaches 

equilibrium. 

The following side reactions produce carbon in the steam reformer: 

2CO Æ C(s) + CO2 (Boudouard coking) 

CO + H2 Æ C(s) + H2O (CO reduction) 

CH4Æ C(s) + 2H2 (Methane cracking) 

The reformer molar steam to carbon ratio is usually between 2-6, depending on the 

feedstock and process conditions. Excess steam is used to prevent coking in the reformer 

tubes. The shift reaction is exothermic and favors low temperatures. Since it does not 

approach completion in the reformer (usually there is 10-15 vol% CO, dry basis, in the 

reformer effluent), further conversion of CO is performed using shift conversion 

catalysts. 

2.6 Catalysts 

Conventional steam reforming catalysts are 10-33 wt% NiO on a mineral support 

(alumina, cement, or magnesia). Reforming catalyst suppliers include, BASF, Dycat 

International, Haldor Topsoe, ICI Katalco and United Catalysts (Leiby, 1994). Heavy 

feedstocks tend to coke the reforming catalyst but promoters (potassium, lanthanum, 

ruthenium, and cerium) may be used to help avoid this problem. These promoters 

increase steam gasification of solid carbon reducing coke formation but they also reduce 

the reforming activity of the catalyst (Leiby, 1994). For feedstocks heavier than naphtha, 

nickel-free catalysts containing primarily strontium, aluminum, and calcium oxides, have 

been successfully tested (Häussinger, et al, 2000). However, the methane content in the 

exiting gas is high, requiring a secondary reformer. 

HTS catalyst has an iron oxide, chromium oxide basis while the major component in the 

LTS catalyst is copper oxide, most often in a mixture with zinc oxide (Häussinger, et al, 

2000). The HTS reactor operates in the temperature range of 300-450 °C while the LTS 

is in the range of 180-270 °C. Often the LTS reactor operates near condensation 

conditions. This catalyst is sensitive to changes in operating conditions. Typical catalyst 

lifetimes for both HTS and LTS catalysts are 3-5 years. 

Sulfur tolerant or “dirty shift” catalysts have also been developed. These catalysts can 

handle larger amounts of sulfur. ICI Katalco makes dirty shift conversion catalysts that 

consist of cobalt and molybdenum oxides. They operate over a temperature range of 

230-500 °C. The controlling factors are the ratio of steam to sulfur in the feed gas and 

the catalyst temperature. 
10

For the methanantion catalyst, nickel oxide is used, often in combination with chromium 

oxide. 

2.7 Reactors 

A wide variety of reformer designs exist and can be used in various process 

configurations. The main components of the reformer furnace include an air/fuel 

combustion system, a radiant heat transfer section, and a convection section. The radiant 

section supplies heat to the catalyst tubes by combustion of the air/fuel mixture and the 

convection section recovers heat by cooling down the flue gases. Reformer furnaces are 

not very efficient and only about half of the heat in the radiant section is absorbed by the 

furnace tubes. 

Generally, the feed gas flows up through the catalyst tubes. However, the reformer 

furnace can be side-, terrace-, top-, or bottom-fired (Johansen, 1992). Top-fired 

reformers are generally suited for larger production units. For small units of < 24 tubes, 

side-fired units are more economical. The terrace-fired reformer, developed in the 1960s, 

is a variation of the side-wall design. Bottom-fired reformers are the least common of the 

four designs (Leiby, 1994). 

The shift converters and methanation reactor are fixed bed reactors. 

2.8 Gas Cleanliness Requirements 

Sulfur compounds are the main poison of reforming catalysts. Even at a concentration of 

0.1 ppm the catalyst can begin to deactivate. To maintain a 3-year catalyst lifetime the 

sulfur concentration in the reformer feed gas should be less than 0.5 ppm (Leiby, 1994). 

Uranium oxide and chromium oxide are used as a promoter in certain reforming catalysts 

resulting in a higher tolerance to sulfur poisoning (Häussinger, 2000). Natural gas usually 

contains only small amounts of sulfur compounds generally in the form of H2S. In most 

SMR plants, prior to the steam reformer, there is a hydrogenator where sulfur compounds 

are converted to H2S. A desulfurizer follows this where the H2S in the feed gas is 

absorbed on a ZnO bed. Any remaining organic sulfur compounds and carbonyl sulfide 

are partially cracked and absorbed on the zinc oxide bed (Häussinger, et al, 2000). If the 

sulfur concentration in the feed gas is greater than 1% then the sulfur must be removed by 

chemical or physical scrubbing. 

The LTS catalyst is also very sensitive to sulfur. Additionally, chloride is a LTS catalyst 

poison. HTS catalyst can tolerate sulfur concentrations up to several hundred ppm, 

although the activity will decline. Other HTS catalyst poisons include phosphorus, 

silicon and unsaturated hydrocarbons in the presence of NOx (Häussinger, 2000). 

Prior to the PSA unit, entrained liquids (water and condensed hydrocarbons) must be 

removed because they will permanently damage the adsorbent, which is a mixture of 

activated carbon and zeolites. Cooling the product and installing a knock out drum with a 

mist eliminator prior to the PSA unit is usually sufficient. 
11

2.9 Status/Technology Developers 

BASF was the first company to develop steam methane reforming in the early 1900s 

(Nirula, 1995). In the mid-1930s to 1950s steam reforming was developed for heavier 

hydrocarbon feedstocks. Thereafter, energy efficiency improvements were made in this 

mature technology by the 1980s. There are many licensors for the various components of 

the steam reforming process. Table 4 is a list of licensors for steam reforming technology 

and Table 5 is a list of current commercial suppliers of PSA units. 

Table 4: Catalytic Steam Reforming Licensors*

Company Headquarters Location 

ABB Lummus Global Inc. Bloomfield, NJ, U.S.A. 

Air Products and Chemicals Allentown, PA, U.S.A. 

BOC Gases Murray Hill, NJ, U.S.A. 

Brown & Root, Inc. Alhambra, CA, U.S.A. 

John Brown Engineers & Constructors Ltd. London, England 

Caloric Gräfelfing, Germany 

Exxon Research & Engineering Co. Florham Park, NJ, U.S.A. 

Haldor Topsoe A/S Copenhagen, Denmark 

Howe-Baker Engineers Tyler, TX, U.S.A. 

International Fuel Cells South Windsor, CT, U.S.A. 

JGC Corporation Tokyo, Japan 

KTI Group BV Zoetermeer, Netherlands 

Koch Process Technology Incorporated Roswell, GA, U.S.A. 

Linde AG Munich, Germany 

Lurgi AG Frankfurt, Germany 

Mahler AGS Stuttgart, Germany 

Messer Griesheim GmbH Krefeld, Germany 

The MW Kellogg Company Houston, TX, U.S.A. 

PETROBRAS – Petróleo Basileiro S/A Ro de Janeiro, Brazil 

Uhde GmbH Dortmund, Germany 

* Source: (Suresh, et al, 2001 and Leiby, 1994) 
12

Table 5: Commercial Suppliers of PSA Units**

Company Headquarters Location 

L’Air Liquide SA Paris, France 

Air Products and Chemicals St. Louis, MO, U.S.A. 

Bergbau AG Herne, Germany 

The BOC Group PLC Windlesham, United Kingdom 

Caloric Analgenbau Munich, Germany 

Howmar United Kingdom 

Linde AG Munich, Germany 

Questor Industries, Inc. Vancouver, BC, Canada 

Toyo Engineering Tokyo, Japan 

UOP, LLC Des Plaines, IL, U.S.A. 

**Source: (Suresh, et al, 2001) 

2.10 Process Performance 

About 65-75% of the CO and steam in the feed stream to the HTS reactor are converted 

to additional hydrogen and CO2 (Leiby, 1994). When a LTS reactor is used to convert 

additional CO to hydrogen about 80-90% of the remaining CO is converted to hydrogen 

resulting in an improvement in the hydrogen yield of about 5% (Leiby, 1994). 

For the PSA unit, the minimum pressure ratio between the feed and purge gas is about 

4:1 and the purge gas pressure is typically between 17-20 psi to obtain a high recovery of 

hydrogen. Hydrogen recovery is usually 85-90% at these conditions and drops to 60-

80% at high purge gas pressures (55-95 psi) (Leiby, 1994). The PSA efficiency is also 

affected by adsorption temperature. Fewer impurities are adsorbed at higher 

temperatures because the equilibrium capacity of the molecular sieves decreases with 

increasing temperature. Additionally, nitrogen is weakly adsorbed onto the adsorbent 

bed in the PSA unit, reducing the hydrogen recovery rate for the same purity. The 

hydrogen recovery may be reduced by as much as 2.5% for a 10-ppm nitrogen 

concentration in the PSA feed stream. 

2.11 Environmental Performance 

Hydrogen itself is a clean burning fuel. However, depending upon the feedstock used, its 

production can generate a considerable amount of CO2. Additionally, steam reformers 

produce NOx from fuel combustion. In California, selective catalytic reduction units are 

used in combination with low NOx burners to meet the state’s strict air regulations 

(Baade, et al, 2001). Controlling the emissions from less hydrogen-rich feedstocks is 

more difficult as the feedstock goes from natural gas to heavy fuel oil, coke, or coal. 

These feedstocks also contain other impurities such as sulfur and heavy metals. 
13

2.12 Cost 

Merchant hydrogen prices vary considerably depending on the volume and form of 

delivery. For large volume users, pipeline is the most economical followed by bulk 

liquid hydrogen delivery. For gaseous hydrogen produced from natural gas at a largescale, central production facility at a pressure of around 400 psi, the plant gate price is 

about $5-$8/GJ. In general, contract prices for delivered hydrogen are not publicly 

available and they vary a great deal depending on the type of delivery, quantity required, 

and delivery distance. SRI states that liquid hydrogen list prices have been around 

$45/GJ, but that the average transaction prices are considerably below this (Suresh, et al, 

2001). A typical price range for large-volume, bulk liquid consumers is $18-$24/GJ 

(Suresh, et al, 2001). 

Generally, as the feedstock goes from natural gas to light hydrocarbons to heavy 

hydrocarbons and then to solid feedstocks, there is an increase in processing difficulty 

and capital costs. Table 6 is a comparison of detailed biomass to hydrogen studies. 

Table 8 gives some costs for hydrogen from natural gas, coal, and biomass from other 

sources containing little or no details. All of these numbers are plant gate costs. 

Table 6: Comparison of Biomass Based Hydrogen Studies 

Study Hamelinck (2001) (a), (b) Larson (1992) (a) Spath (2000) (a) 

Cost year 2001 1991 2000 

Biomass feed rate 

(BD tonne/day) 

1,920 1,650 about 300-1,500 

(3 plant sizes) 

Biomass cost $2/GJ 

$34/dry tonne 

$2/GJ 

$34/dry tonne 

$2.7/GJ 

$46.2/dry tonne 

Electricity selling price $0.03/kWh $0.05/kWh 

(purchased) 

$0.05/kWh (purchased) 

Net power (MW) ranges from no excess to 

about 85 MW (@ 91% 

production cap) 

no excess no excess (excess 

steam is produced and 

sold) 

H2 produced (tonne/day) 91 – 184 (@ 91% 

production cap) 

134 - 188 23 – 114 

Price of H2 (HHV basis, 

plant gate) 

$8-11/GJ $7-12/GJ $14-21/GJ 

(for 15% IRR; 

gasification only) 

Level of detail Very detailed - Gives costs 

of individual equipment, 

operating costs, and other 

economic parameters and 

assumptions 

Somewhat 

detailed – Gives 

capital and 

operating costs 

Somewhat detailed - 

Gives total installed 

capital costs, and other 

economic parameters 

and assumptions 

Notes: (a) Examined direct & indirect gasifiers and atmospheric and pressurized gasifiers. 

(b) Several configurations examined (varied type of gas cleaning, and H2 separation) 
14

Table 7: Hydrogen Cost Data From other Sources 

Resource IEA (1999) IEA 

(2001) 

McKinley, 

et al (1990) 

British 

Government 

Panel (1999) 

Grégoire 

Padró 

(1999) 

Amick 

(2003) 

Natural 

gas 

$5.6/GJ 

(nat gas @ 

$3/GJ) 

$7/GJ 

(w/CO2 seq) 

$5.4-

7.5/GJ 

$7/GJ 

(1987$) 

$5.7/GJ 

(nat gas @ 

$3/GJ) 

$7/GJ 

(w/CO2 seq) 

$5.97-

7.46/GJ 

----- 

Coal $10/GJ 

(coal @ $2/GJ) 

$13/GJ 

(w/CO2 seq) 

$10-12/GJ $15/GJ 

(1987$) 

$10.3/GJ 

(coal @ $2/GJ) 

$13.2/GJ 

(w/CO2 seq) 

$9.87-

19.3/GJ 

$9-10/GJ 

Biomass ----- $12-13/GJ $7/GJ 

(Biomass 

@ $15/ton - 

1987$) 

----- $8.69-

17.1/GJ 

----- 

2.13 R&D Needs 

Although steam reforming has been around for many years, improvements in reforming 

and shift catalyst continue to be made. Additionally, more efficient engineering designs 

have aided in cost reductions. 

2.14 References 

Amick, P. Geosits, R., Kramer, S. (2003). 225th ACS National Meeting. New Orleans, 

Louisiana. March 23-27. Fuel Chemistry Preprints. 48 (1) pg. 199-204. 

Baade, W.F.; Parekh, U.N.; Raman, V.S. (2001). "Hydrogen." Kirk-Othmer 

Encyclopedia of Chemical Technology, John Wiley & Sons, Inc. 

British Government Panel on Sustainable Development. July 1999. 

Grégoire Padró, C.E.; Putsche, V. (1999). Survey of the Economics of Hydrogen 

Technologies. 57 pp.; NREL Report No. TP-570-27079. 

Häussinger, P.; Lohmüller, R.; Watson, A.M. (2000). "Hydrogen." Ullmann's 

Encyclopedia of Industrial. Wiley-VCH Verlag GmbH & Co.KGaA. 

IEA Greenhouse Gas R&D Programme (1999). Hydrogen – Today and Tomorrow. 

IEA Bioenergy Task 33: Thermal Gasification of Biomass, Spring 2001, Task Meeting. 
15

Johansen, T.; Raghuraman, K.S.; Hackett, L.A. (1992). “Trends in hydrogen plant 

design”. Hydrocarbon Processing. 

Leiby, S.M. (1994). Options for Refinery Hydrogen. SRI Report No. 212. Menlo Park, 

CA. 

McKinley, K.R., Browne, S.H., Neill, D.R., Seki, A., Takahashi, P.K. (1990). 

“Hydrogen Fuel from Renewable Resouorces”. Energy Sources. Vol 12. pp. 105-110. 

Nirula, S.C. (1995). Synthesis Gas. SRI Report No. 148A. Menlo Park, CA. 

Suresh, B.; Gubler, R.; Sasano, T. (2001). "Hydrogen." Chemical Economics Handbook 

Product Review, SRI International, Menlo Park, CA. Reprot number 743.5000. 

Spath, P.L.; Lane, J.M; Mann, M.K; Amos, W.A. (April 2000). Update of Hydrogen 

from Biomass - Determination of the Delivered Cost of Hydrogen. NREL Milestone 

report. 

Wender, I. (1996). "Reactions of synthesis gas." Fuel Processing Technology 48(3): 

189-297. 
16

3 Ammonia 

3.1 Summary 

Ammonia is the second largest synthetic chemical product. Most of the ammonia 

produced is used to make fertilizers in the form of urea or ammonium salts. The basic 

process steps for industrial production are: synthesis gas production, gas conditioning, 

compression, and ammonia synthesis. The goal of the synthesis gas production and gas 

conditioning steps is to provide a pure H2 stream for input into the ammonia converter. 

Ammonia synthesis is achieved in catalytic converters at pressures ranging from 15-35 

MPa at a minimum temperature of 430-480°C. Ammonia synthesis occurs via the 

following exothermic chemical reaction: N2 + 3H2 → 2 NH3 and is typically performed 

over promoted Fe catalysts. The optimum H2/N2 ratio is near 2 at high space velocities 

and approaches 3 at low space velocities as equilibrium becomes more dominant. 

Commercial ammonia synthesis catalysts are basically the same Fe-promoted catalysts 

first developed by Mittasch in 1910. However, the search for catalysts with high activity 

and long lifetimes at lower synthesis temperatures and lower pressures continues. Other 

catalysts are being developed such as promoted Ru on high surface area graphite 

supports. These recently developed commercial catalysts offer the possibility of greatly 

reducing synthesis temperatures and pressures with improved activity at high ammonia 

concentrations. 

The basic design of an ammonia synthesis reactor is a pressure vessel with sections for 

catalyst beds and heat exchangers. Ammonia converters are classified by flow type 

(radial, axial, or cross flow) and cooling method (quench or indirect) used. Removing the 

heat generated from the exothermic synthesis reaction to maintain control of the reaction 

temperature is the main design challenge. 

Gas phase impurities can be separated into temporary and irreversible catalyst poisons. 

Oxygenated compounds like CO, CO2, H2O, and O2 are temporary, reversible ammonia 

catalyst poisons at low process temperatures. Removal of these oxygenates from the feed 

gas stream followed by re-reduction of the catalyst with pure H2 and N2 will restore 

catalyst activity. Sulfur, Chlorine, and As, P, Sb are irreversible poisons. 

In the early days the ammonia industry was based on coal as the feedstock but as cheap 

petroleum feedstocks became available the industry moved toward natural gas. The cost 

of ammonia is highly dependent on the feedstock price. In general, the price of ammonia 

has fluctuated between $100-$250/tonne since 1986. Improved catalysts and process 

optimization are part of the continuing efforts to improve the economics of ammonia 

production by reducing process energy consumption. 
17

Section Contents 

3.1 Summary ........................................................................................................................ 16

3.2 Introduction ................................................................................................................... 18

3.3 Technology Description................................................................................................. 19

3.4 Chemistry....................................................................................................................... 20

3.5 Catalysts........................................................................................................................ 21

3.6 Reactors......................................................................................................................... 23

3.7 Gas Cleanliness Requirements...................................................................................... 25

3.8 Status/Technology Developers ...................................................................................... 26

3.9 Environmental Performance ......................................................................................... 27

3.10 Cost.............................................................................................................................. 27

3.11 R&D Needs.................................................................................................................. 28

3.12 References.................................................................................................................... 28

List of Figures 

Figure 3: Generic Ammonia Synthesis Process Flow Diagram............................................ 19 

Figure 4: Various Ammonia Converter Designs................................................................... 24 

List of Tables 

Table 8: Summary of Commercial Ammonia Synthesis Catalyst Compositions.................... 23 

Table 9: Summary of Ammonia Synthesis Catalyst Poisons.................................................. 26 

Table 10: Comparison of Ammonia Economics.................................................................... 27 
18

3.2 Introduction 

Ammonia is the second largest synthetic chemical product. In 1996, worldwide ammonia 

production totaled 143 million tonnes (Appl 2003). Most of the ammonia produced is 

used to make fertilizers in the form of urea or ammonium salts (nitrate, phosphate, and 

sulfate). A small fraction of ammonia is used for the manufacture of organic chemical 

feedstocks for the plastics industry; polyamides, caprolactam, and others, and for the 

production of explosives (hydrazine, nitriles, etc.). Ammonia is also converted to nitric 

acid and cyanides. Ammonia is manufactured from nitrogen, fixed from the atmosphere, 

and hydrogen, produced mainly by steam methane reforming, in the catalytic process 

developed in the early 1900s by Fritz Haber and Carl Bosch using a promoted iron 

catalyst discovered by Alwin Mittasch (Satterfield 1991). The onset of World War I 

brought concerns about diminished worldwide nitrate supplies caused by an increase in 

the production of explosives. This led to the first industrial plant constructed at Oppau, 

Germany, and commissioned in 1913. About 50% of the hydrogen produced from 

syngas processes is used for ammonia production (Wender 1996). 

The vast amount of literature on ammonia synthesis over the past 100 years details 

process improvements, reactor engineering, catalyst developments, and mechanistic 

interpretation of the conversion of N2 and H2 to NH3 over Fe catalysts at high 

temperature and high pressure. Several reviews on ammonia production were consulted 

in writing this technology description (Kirth-Othmer)(Appl 2003; Jennings and Ward 

1989; Satterfield 1991) and individual references to previous work can be found therein. 

For example, the “Ammonia” article in Ullmann’s Encyclopedia of Industrial Chemistry 

contains over 1000 cited references (Appl 2003). Numerous books and entire journals 

such as Nitrogen and Ammonia Plant Safety are can be found that are dedicated to all 

aspects of ammonia synthesis. 

Commercial ammonia synthesis catalysts are basically the same Fe promoted catalysts 

first developed by Mittasch. The search for catalysts with high activity and long lifetimes 

at lower synthesis temperatures and lower pressures continues. Recent developments in 

ammonia synthesis catalysts include Ru-based formulations that are showing promise for 

commercial applications. 

Improved catalysts and process optimization are part of the continuing efforts to improve 

the economics of ammonia production by reducing process energy consumption. Energy 

consumption in ammonia plants has decreased nearly an order of magnitude since the 

early plants; from ~100 GJ/tonne NH3 to ~30 GJ/tonne NH3 (Jennings and Ward 1989). 

Part of this is a result of improved efficiencies in syngas (hydrogen) production by 

switching fuels, from coal to natural gas. Improved reactor design and optimized heat 

integration in the ammonia plant also contributes significantly to reducing energy 

consumption. One objective is to optimize heat recovery and minimize the capital and 

operating costs of the ammonia synthesis loop. 
19

3.3 Technology Description 

The basic process steps for industrial ammonia production are synthesis gas production, 

gas conditioning, compression, and ammonia synthesis. The goal of the synthesis gas 

production and gas conditioning operations is to provide a pure H2 stream for input into 

the ammonia converter. This requires a number of unit operations including feedstock 

pretreatment, syngas generation, CO conversion, and gas purification. These individual 

operations are summarized in the generic block flow diagram in Figure 3. 

Figure 3: Generic Ammonia Synthesis Process Flow Diagram 

For natural gas plants, feedstock pretreatment is essentially desulfurization prior to steam 

methane reforming. Feedstock preparation and handling is more complicated for solid 

fuel based systems. Secondary reformers are also used to further decrease the 

hydrocarbon content in the syngas. Following the reforming section, the syngas is 

cleaned and conditioned. Syngas cleaning is more involved for solid fuel based 

processes. Conditioning the cleaned syngas is required to increase the yield and purity of 

the hydrogen that is produced. CO is converted to H2 via the water gas shift reaction in 

high and low temperature shift reactors. Water is removed by cooling the syngas and 

CO2 is removed in an absorber-stripper process. Residual CO and CO2 in the product H2

is reduced to lower levels by what amounts to the reverse steam methane reforming 

process. A fraction the H2 is consumed in a methanation reactor to convert residual CO 

and CO2 to CH4. Purified H2 is then compressed and fed to the ammonia synthesis loop. 

Another configuration would be to replace the following process steps: low temperature 

shift, CO2 removal, methanation and possibly the secondary reformer with pressure 

Gas 

Pretreatment

Gas 

Pretreatment

Primary

Reforming

Gas

Steam

Secondary 

Reforming

Air

CO Shift

Conversion

Compression

H2 Recycle

Methanation CO2 Removal

CO2

Ammonia 

Synthesis

Synthesis Gas 

Recycle

Refrigeration Product

Purge

Hydrodesulfurization

ZnO adsorption

Saturation

SMR Partial Oxidation High and Low

Temperature

MEA/DEA

Rectisol (CH3OH)

Selexol

K2CO3
20

swing adsorption. It is also possible to replace the reforming section with an autothermal 

reformer. However, currently the majority of ammonia plants follow the process scheme 

shown in Figure 3. 

Ammonia synthesis is achieved in catalytic converters at pressures ranging from 15-35 

MPa at a minimum temperature of 430-480°C. A maximum temperature of ~550°C 

results from the balance between increasing reaction kinetics and decreasing the 

equilibrium NH3 concentration as the temperature increases. Thermodynamics suggests 

that low process temperatures are favorable, however, the kinetics of the reactions 

dictates that high temperatures are required. The most effective catalysts from a process 

perspective are those that have the highest rate of conversion at the lowest temperatures 

(Jennings and Ward 1989). With effective gas cleanup and conditioning, catalyst 

lifetimes of 5-8 years can be achieved in commercial practice and up to 14 years in 

certain cases. Ammonia formation increases as the process pressure is increased. The 

optimum H2/N2 ratio is near 2 at high space velocities and approaches 3 at low space 

velocities as equilibrium becomes more dominant. Space velocities in commercial 

ammonia synthesis vary from 12,000/hr at 15 MPa to 35,000/hr at 80MPa (Appl 2003). 

Per-pass conversion on the order of 10-35% is typically achieved. Ammonia is recovered 

from the synthesis loop by cooling the synthesis gas at process pressures to condense the 

ammonia. The liquid ammonia is separated from the gas, which is recycled back through 

the converter. Earlier plant designs used air or water-cooling for ammonia recovery. 

Modern synthesis plants use refrigeration to condense out the ammonia. 

The ammonia recovery is not extremely efficient so the recycled gas typically contains 

4% NH3 plus any inert gases (Ar, He, CH4, etc.) that may be in the process stream. 

Purging some of the gas in the recycle loop before it is recycled minimizes inert gas 

concentrations. This purged gas also contains some product NH3 that is recovered, 

unconverted N2 and H2, plus the inert gases. The purged H2 is recovered and recycled 

back to the converter or used as a fuel in the process. Some of the inert gases also end up 

dissolved in the condensed NH3 product. Flashing the liquid ammonia in a pressure 

letdown step releases any dissolved gases that are separated from the product. 

Clearly, the ammonia synthesis process consists of many complex unit operations apart 

from the actual synthesis loop. The way in which these process components are 

combined with respect to mass and energy flow has a major influence on efficiency and 

reliability. Apart from the feedstock, many of the differences between various 

commercial ammonia processes lie in the way in which the process elements are 

integrated. 

3.4 Chemistry 

Ammonia synthesis occurs via the following exothermic chemical reaction: 

N2 + 3H2 → 2 NH3 ∆H773K = -109 kJ/mol; ∆H298K = -46.22 kJ/mol 
21

Since the reaction is exothermic, maximum conversion at equilibrium occurs at high 

pressure and low temperature. Early attempts at ammonia production were hampered by 

a lack of understanding of thermodynamic equilibrium and the temperatures and 

pressures chosen for ammonia synthesis were unfavorable for reaction. 

Ammonia synthesis is typically performed over promoted Fe catalysts. Numerous 

mechanistic studies based on the measurement of surface adsorbed species have 

determined that the rate-limiting step is the dissociative chemisorption of N2 on the 

catalyst surface to form N atoms. Surface concentration of N atoms is governed by the 

equilibrium with gas phase H2 and NH3. 

The following rate expression for ammonia synthesis was developed by Temkin and 

Pyzhev in 1940 and has been widely applied to ammonia converter design and 

development (Jennings and Ward 1989) and references therein): 

1 , 0 3

2

2 2

3

1

2

3

3

2

2 + = > 













 − 













 = α β α β

α β

H

NH

NH

H

N

P

P

k

P

P

r k P

where the first term is the rate of ammonia formation and the second term is the rate of 

ammonia decomposition. This rate expression is valid over a very wide pressure range 

with fugacities substituted for pressures at very high pressure. The value for α ranges 

between 0.4-0.75 for commercial processes where ammonia is recycled back to the 

reactor. Clearly, this expression is not valid for zero ammonia concentrations that are 

often encountered in the laboratory. For process using only pure reactants without 

ammonia recycle, the following rate equation applies: 

0.5 1

2 2 = = − α α α

H PN r kP

3.5 Catalysts 

Mittasch screened more than 2000 catalyst compositions before developing a promoted 

Fe catalyst that is, in principle, still commercially used today. Catalysts usually contain 

multiple promoters to improve the effectiveness and stability of the catalysts for long 

periods of time. 

Ammonia synthesis catalysts are made by cooling, casting, crushing, and sieving a fused 

melt of magnetite ore (Fe3O4) and promoter precursors. Promoters include Al2O3, K2O, 

CaO, MgO, etc. One of the functions of the promoters is to minimize Fe sintering during 

reduction and reaction. K2O is a structural promoter. Typical ammonia synthesis 

catalysts contain: 2.5-4% Al2O3, 0.5-1.2% K2O, 2.0-3.5% CaO, 0-1.0% MgO, and 0.2-

0.5% SiO2 (naturally occurring in the magnetite ore). Commercial catalyst formulations 

are listed in Table 8. 
22

Finely divide iron is pyrophoric so Fe catalysts are reduced to metallic form in situ in 

ammonia converters. Porosity is developed as the catalysts are reduced. During 

reduction, oxygen is removed from the magnetite crystal lattice without shrinkage. 

Metallic Fe is formed with essentially the same porous structure as the magnetite 

precursor (Jennings and Ward 1989). The production and preservation of this highly 

porous structure during reduction of the ammonia synthesis catalysts leads to highly 

active catalysts. The additions of structural promoters like Al2O3, MgO, and Cr2O3

facilitates the formation of highly porous metallic iron. Surface areas of freshly reduced 

promoted iron ammonia synthesis catalysts can be as high as 15-20 m2

/g. 

Catalyst activity is not directly associated with Fe surface area but is also related to the 

complex interactions of the promoters. Alkali metal promoters in ammonia synthesis 

catalysts are needed to attain high activity. Potassium is the most cost effective alkali 

promoter. Cs and Rb are effective promoters but more costly than K. Li and Na are poor 

promoters and not used commercially. Potassium is thought to interact with the Fe 

crystallites by increasing the dissociative sticking probability of N2 on the Fe sites and 

increasing catalytic activity. Dissociative adsorption of H2 on the Fe sites occurs rapidly 

at relatively low temperatures. Equilibrium is, therefore, attained for H2

adsorption/desorption at ammonia synthesis temperatures. 

Other catalysts are being developed such as promoted Ru on high surface area graphite 

supports. These recently commercial catalysts offer the possibility of greatly reducing 

synthesis temperatures and pressures with improved activity at high ammonia 

concentrations. The Kellogg Advanced Ammonia Process is based on a Ru catalyst that 

is claimed to be 40% more active than Fe catalysts. 
23

Table 8: Summary of Commercial Ammonia Synthesis Catalyst Compositions 

Catalyst Composition (wt%) Operating Ranges 

Vendor Name Fe oxides# K2O Al2O3 CaO SiO2 Other T (°C) P (bar) 

Haldor KM1 91-95 5-9 

Topsoe 

KM1R* 89-93 7-11 

340-550 88-588 

35-4 93.9 0.7 3.0 2.4 

35-8* 92.1 0.9 3.8 3.2 

350-530 100-600 

S6-10 93.9 0.7 3.0 2.4 

S6-10R* 92.1 0.9 3.8 3.2 

BASF 

catalyst 

 

74-1 95.8 0.6 2.1 1.5 

Katalco 

(ICI) 

74-1R* 94.7 0.8 2.6 1.9 

≥60 

Sud AS-4 92 promoters 

Chemie 

As-4R* 78%Fe 10% 

FeO 

promoters 

* pre-reduced catalysts 

# For unreduced catalysts, Fe is in the form of magnetite (Fe3O4). In pre-reduced catalysts the Fe 

is present as metallic Fe and FeO. 

3.6 Reactors 

The basic design of an ammonia synthesis reactor is a pressure vessel with sections for 

catalyst beds and heat exchangers. Ammonia converters are classified by flow type 

(radial, axial, or cross flow) and cooling method (quench or indirect) used. 

Axial flow reactors are essentially top-to-bottom flow reactors. The design is 

comparatively simple, however, a fairly large pressure drop develops across the catalyst 

bed. A radial flow configuration feeds gas into an annular region between the reactor 

wall and the outer surface of the catalyst bed. Gas flows through the bed and exits out a 

central collection tube. This design minimizes pressure drop across a shallow bed with a 

large surface area. Radial flow converters tend to be tall vessels with relatively small 

diameter. The cross-flow configuration has a similar principle in that gas is introduced 

along one side of the reactor and is collected radially across the reactor by a collector on 

the other side. 

Removing the heat generated from the exothermic synthesis reaction to maintain control 

of the reaction temperature is another design challenge. Quench converters are designed 

to introduce cool reactant gas at various points along the length of the catalyst bed. 

Interbed heat exchangers can also be used to remove heat at specific intervals along the 

bed, effectively separating the bed into multiple synthesis zones, or continuously along 

the bed with cooling tubes. These indirectly cooled designs allow for efficient reaction 
24

heat recovery that can be used in other parts of the process. Reactant gas can be 

circulated through the heat exchangers to preheat the ammonia synthesis gas or water is 

used to produce steam. 

A number of ammonia converters are commercially available. Seventeen different 

designs for ammonia converters are described by Zardi (Zardi 1982). The article in 

Ullmann’s also contains figures and descriptions of several different ammonia converters 

(Appl 2003). Many designs are based on quench converters that have multiple catalyst 

beds in series and cold gas is introduced between the beds for temperature control. Early 

designs based on axial flow were easier to build but not as efficient as the more complex 

radial flow converters that have more recently been designed. A small subset of 

ammonia converters are depicted in Figure 4 and described below.

Figure 4: Various Ammonia Converter Designs 
25

3.6.1 Topsoe radial-flow converter 

One design uses 2 radial beds with quench gas injection between them. A similar radial 

flow design uses an inter-bed heat exchanger in the first catalyst bed. Cold ammonia 

synthesis gas is introduced from the bottom of the converter through the second catalyst 

bed then through the heat exchanger in the first catalyst bed. The cold gas flow through 

the second bed also provides indirect heat exchange. An additional heat exchanger is 

located at the bottom of the reactor to cool the reacted gases. 

3.6.2 Kellogg 4-bed axial flow quench converter 

This reactor consists of 4 catalyst beds held on separate grids. Quench gas is introduced 

in the spaces between the beds. A heat exchanger is located at the top of the vessel. 

3.6.3 Kellogg horizontal converter 

This reactor is a cross-flow converter design where gas flows through the catalyst bed 

perpendicular to the axis of the vessel. It is available in both quench and indirectly 

cooled versions. 

3.6.4 Tube Cooled Converter 

There are many variations in this design. Figure 4 shows a generic reactor. 

3.6.5 ICI Lozenge Quench Converter 

This is an additional reactor which is not shown in Figure 4. An axial flow, continuous 

catalyst bed is divided by lozenge distributors for quench gas addition. 

3.7 Gas Cleanliness Requirements 

In a carefully controlled ammonia synthesis processes, 10-year catalyst lifetimes are not 

uncommon. Loss of activity, however, does still occur by common deactivation 

mechanisms. Deactivation by sintering is a slow, gradual process that leads to loss of 

active area by Fe crystallization and oxidation. The activity of ammonia synthesis 

catalysts can also be reduced by trace impurities in the feed gas stream or solid impurities 

introduced during catalyst manufacture. Careful manufacturing processes are required to 

avoid the solid catalyst poisons, although, some are unavoidable because they are present 

as impurities in the natural magnetite starting material. Extensive purification of H2 and 

N2 is required to minimize gas phase catalyst poisons. 

Gas phase impurities can be separated into temporary and irreversible catalyst poisons. 

Oxygenated compounds like CO, CO2, H2O, and O2 are temporary, reversible ammonia 

catalyst poisons at low process temperatures. Removal of these oxygenates from the feed 

gas stream followed by re-reduction of the catalyst with pure H2 and N2 will restore 
26

catalyst activity. A typical maximum specification allowed for all oxygen containing 

compounds in the ammonia synthesis reactor feed is 10 ppm (total oxygen compounds). 

As temperatures increase above 520°C water vapor can become an irreversible poison as 

Fe crystallites form. 

Very low levels of gaseous sulfur and chlorine irreversibly poison catalysts. The sulfur is 

removed from the process gas stream before the primary reformer by a 

hydrodesulfurization unit containing Co-Mo catalysts followed by a fixed bed ZnO 

absorber. Syngas from coal gasification can contain high levels of H2S and COS that 

must be removed. 

Chlorine attacks the alkali promoters to form volatile alkali chlorides that can vaporize at 

process temperatures. Chlorine also attacks the active metallic Fe sites converting them 

to Fe-chlorides. Trace levels of arsenic and phosphorous are also strong catalyst poisons. 

Table 9 summarizes the various catalyst poisons. 

Table 9: Summary of Ammonia Synthesis Catalyst Poisons 

Poison Type Maximum 

concentration (ppm) 

H2O Rev 200 

CO Rev 200 

CO2 Rev 100 

O2 Rev 100 

Sulfur (H2S) Irrev 0.1 

Chlorine irrev 0.1 

As, P, Sb Irrev --- 

3.8 Status/Technology Developers 

In the early days the entire ammonia industry was based on coal as the feedstock. Today 

coal or coke (including coke oven gas) is no longer of major economic importance. In 

1990, for example, only 13.5 % of the world ammonia capacity was based on this raw 

material. Apart from a few plants operating in India and South Africa, today, the 

majority of coal-based ammonia plants are found in China. As cheap petroleum 

feedstocks became available the industry moved toward natural gas as the main 

feedstock. 

More than 20 commercial ammonia synthesis processes are described in the article on 

“Ammonia” in Ullmann’s Encyclopedia of Industrial Chemistry (Appl 2003). 

Commercial vendors include: 
27

ICI – (http://www.synetix.com/ammonia/index.htm) 

Linde (Ilg and Kandziora 1997) - (http://www.linde-processengineering.com/en/p0001/p0017/p0020/p0020.jsp), 

Kellogg, Brown, and Root - (http://www.mwkl.co.uk/pdf/Ammonia_brochure_1.pdf and 

http://www.mwkl.co.uk/pdf/KAAPplus.pdf) 

Haldor Topsoe (http://www.topsoe.com/tempfiles/267.asp) 

Ammonia Casale (http://www.casale.ch/ammonia/) 

Krupp-Uhde - (http://www.uhde.biz/kompetenzen/technologien/duengemittel.en.html) 

3.9 Environmental Performance 

There are safety aspects in handling and storing liquid ammonia because it is toxic. The 

release of CO2 is an environmental concern related to ammonia production. Often times 

the CO2 is captured and sold, otherwise it is vented to the atmosphere. 

3.10 Cost 

The cost of ammonia is highly dependent on the feedstock price. In general, the price of 

ammonia has fluctuated between $100-$250/tonne since 1986 (Appl 2003). It is 

interesting to note that in 1969 the majority of the ammonia was produced in North 

America and Europe (54% of total worldwide production) and now Asia produces the 

largest fraction of ammonia (38% of the total in 1996). Table 10 gives the cost of 

ammonia from a few sources for different feedstocks. 

Table 10: Comparison of Ammonia Economics 

Study Dietz et al 1978 Appl 2003 Eggeman 

2001 

Cost year 1981 1980 1980 1980 1996 1990 

Feedstock Brava cane Natural 

gas 

Resid Coal Natural 

gas 

Vacuum 

resid 

Coal Natural gas 

Feedstock 

cost 

$9.77/dry 

ton 

na na na $2.8/GJ $2.0/GJ $1.8/GJ $1.9/GJ 

Ammonia 

(tonne/day) 

91-363 900 900 900 na na na 1,000 

Price of 

ammonia 

($/tonne) 

$213-314 $215 $271 $314 $206 $248 $362 $161 

Level of 

detail 

Some details Minimal details Minimal 

details 
28

3.11 R&D Needs 

One way to improve the economics of ammonia synthesis is by improved heat integration 

and more efficient H2 production. Additionally, advances in better low temperature 

ammonia synthesis catalysts with high activity and long lifetime need to be made. 

3.12 References 

Appl, M. (2003). "Ammonia." Ullmann's Encyclopedia of Industrial Chemistry Release 

2003, 6th Edition, Wiley-VCH Verlag GmbH & Co.KGaA. 

Dietz, C.K., Lutz, I.H., Roman, A. (1978) Ammonia Production from Brava Cane. 

Energy From Biomass and Wastes Symposium, August 14-18, Washington, D.C., 575-

604. 

Eggeman, T. (2001). "Ammonia." Kirk-Othmer Encyclopedia of Chemical Technology, 

John Wiley & Sons, Inc. 

Ilg, J., and Kandziora, B. (1997). "Linde ammonia concept." Ammonia Plant Safety & 

Related Facilities, 37, 341-352. 

Jennings, J. R., and Ward, S. A. (1989). "Ammonia Synthesis." Catalyst Handbook, M. 

V. Twigg, ed., Wolfe Publishing, Ltd., London, England, 384-440. 

Lauriente, D.H. (2001). "Ammonia." Chemical Economics Handbook Marketing 

Research Report, SRI International, Menlo Park, CA. Report number 756.6000. 

Satterfield, C. N. (1991). Heterogeneous Catalysis in Industrial Practice, Krieger 

Publishing Company, Malabar, FL. 

Zardi, U. (1982). "Review These Developments in Ammonia and Methanol Reactors." 

Hydrocarbon Processing, 61(8), 129-133. 
29

4 Methanol and Methanol Derivative (including 

formaldehyde, acetic acid, MTG, MTO, MOGD, TIGAS, and DME)

4.1 Summary 

Currently, the majority of methanol is synthesized from syngas that is produced via steam 

reforming of natural gas. Catalytic methanol synthesis from syngas is a classic hightemperature, high-pressure, exothermic, equilibrium limited synthesis reaction with 

overall conversion efficiency of over 99%. Removing the large excess heat of reaction 

and overcoming the thermodynamic constraint are challenges to overcome in commercial 

methanol synthesis. Numerous methanol converter designs have been commercialized 

over the years and these can be roughly separated into two categories - adiabatic or 

isothermal reactors. Examples of commercially available methanol converters are 

described. 

The first high-temperature, high-pressure methanol synthesis catalysts were ZnO/Cr2O3

and were operated at 350°C and 250-350 bar. Over the years, as gas purification 

technologies improved (i.e., removal of impurities such as sulfur, chlorine, and metals) 

interest in the easily poisoned Cu catalysts for methanol synthesis was renewed. In 1966, 

ICI introduced a new, more active Cu/ZnO/Al2O3 catalyst that began a new generation of 

methanol production by a low temperature (220-275°C), low pressure process (50-100 

bar). Under normal commercial operating conditions, CuO/ZnO methanol catalysts have 

quite long lifetimes, up to 2-5 years. Copper catalysts are extremely sensitive to siteblocking poisons such as reduced sulfur. To retain the long-term activity of Cu catalysts, 

it has been found empirically that the gas-phase sulfur concentration needs to be kept 

below 1 ppm and preferably below 0.1 ppm. If reactor temperatures are not properly 

controlled, the highly exothermic methanol synthesis reactions can also rapidly cause 

catalyst deactivation by sintering of the Cu crystallites. Reactor temperatures are 

maintained below 300°C to minimize sintering. The presence of Cl in syngas has also 

been correlated with a greatly enhanced rate of sintering of copper crystallites most likely 

due to the formation of volatile copper chloride compounds. The limits on HCl content 

to avoid catalyst poisoning are on the order of 1 ppb. 

Methanol is a commodity chemical, one of the top ten chemicals produced globally. A 

summary of the literature regarding the economics of methanol produced from syngas is 

presented. The many uses of methanol are also described. Methanol can be used directly 

or blended with other petroleum products as a clean burning transportation fuel. 

Methanol is also an important chemical intermediate used to produce a number of 

chemicals, including: formaldehyde, dimethyl ether, methyl tert-butyl ether, acetic acid, 

and olefins, to name a few. A description of the processes associated with these 

secondary products is presented with a summary of the available economics for each 

process. 
30

Section Contents 

4.1 Summary ........................................................................................................................ 29

4.2 Introduction ................................................................................................................... 31

4.3 Technology Description................................................................................................. 31

4.4 Reactors......................................................................................................................... 32

4.5 Chemistry....................................................................................................................... 34

4.6 Catalysts........................................................................................................................ 35

4.7 Gas Cleanliness Requirements...................................................................................... 36

4.8 Status/Technology Developers ...................................................................................... 38

4.9 Process Performance..................................................................................................... 39

4.10 Uses ............................................................................................................................. 40

4.11 Environmental Performance........................................................................................ 47

4.12 Cost of Methanol ......................................................................................................... 47

4.13 R&D Needs.................................................................................................................. 51

4.14 References.................................................................................................................... 52

List of Figures 

Figure 5: Simplified Methanol Synthesis Process Flow Diagram......................................... 31 

List of Tables 

Table 11: Commercial Methanol Synthesis Catalyst Formulations...................................... 36 

Table 12: MeOH Gaseous Contaminant Constraints............................................................ 38 

Table 13: Methanol Technology Suppliers............................................................................ 39 

Table 14: Conversions for Methanol, Methanol/DME, and DME ........................................ 45 

Table 15: Comparison of Coal Based Methanol Economic Studies...................................... 48 

Table 16: Comparison of Biomass Based Methanol Studies................................................. 49 

Table 17: Comparison of Formaldehyde Production Costs.................................................. 50 

Table 18: Comparison of Coal Based MTG Economic Studies............................................. 51 
31

4.2 Introduction 

Methanol synthesis actually began in the 1800s with the isolation of “wood” alcohol from 

the dry distillation (pyrolysis) of wood. Research and development efforts at the 

beginning of the 20th century involving the conversion of syngas to liquid fuels and 

chemicals led to the discovery of a methanol synthesis process concurrently with the 

development of the Fischer-Tropsch synthesis. In fact, methanol is a byproduct of 

Fischer-Tropsch synthesis when alkali metal promoted catalysts are used. Methanol 

synthesis is now a well-developed commercial catalytic process with high activity and 

very high selectivity (>99%). For economic reasons, methanol is almost exclusively 

produced via reforming of natural gas (90% of the worldwide methanol, Davenport 

2002). However, a variety of feedstocks other than natural gas can be used. 

The long time interest in methanol is due to its potential fuel and chemical uses. In 

particular, methanol can be used directly or blended with other petroleum products as a 

clean burning transportation fuel. Methanol is also an important chemical intermediate 

used to produce: formaldehyde, dimethyl ether (DME), methyl tert-butyl ether (MTBE), 

acetic acid, olefins, methyl amines, and methyl halides, to name a few. 

4.3 Technology Description 

Currently, the majority of methanol is synthesized from syngas that is produced via steam 

reforming of natural gas. It can also be reformed using autothermal reforming (ATR) or 

a combination of steam methane reforming (SMR) and ATR. Once the natural gas is 

reformed the resulting synthesis gas is fed to a reactor vessel in the presence of a catalyst 

to produce methanol and water vapor. This crude methanol, which usually contains up to 

18% water, plus ethanol, higher alcohols, ketones, and ethers, is fed to a distillation plant 

that consists of a unit that removes the volatiles and a unit that removes the water and 

higher alcohols. The unreacted syngas is recirculated back to the methanol converter resulting in 

an overall conversion efficiency of 99%. A generic methanol synthesis process flow diagram is 

shown in Figure 5. 

Figure 5: Simplified Methanol Synthesis Process Flow Diagram 

Desulphurization Natural

Gas Steam Reforming

Steam, O2

syngas (CO/CO2/H2)

Compressor Methanol

Converter

Cooling

and

Distillation

Methanol

syngas recycle loop

Purge

Gas
32

As is the case with Fischer-Tropsch synthesis, one of the challenges associated with 

commercial methanol synthesis is removing the large excess heat of reaction. Methanol 

synthesis catalyst activity increases at higher temperatures but so does the chance for 

competing side reactions (Supp and Quinkler 1984). Byproducts of methanol formation 

are: CH4, DME, methyl formate, higher alcohols and acetone. Catalyst lifetimes are also 

reduced by continuous high temperature operation and typically process temperatures are 

maintained below 300°C to minimize catalyst sintering. 

Overcoming the thermodynamic constraint is another challenge in commercial methanol 

synthesis. The maximum per-pass conversion efficiency of syngas to methanol is limited 

to about 25% (Wender 1996). Higher conversion efficiencies per-pass can be realized at 

lower temperatures where the methanol equilibrium is shifted towards products; however, 

catalyst activities generally decrease as the temperature is lowered. Removing the 

methanol as it is produced is another strategy used for overcoming the thermodynamic 

limitations and improving the per pas conversion process efficiencies. Methanol is either 

physically removed (condensed out or physisorbed onto a solid) or converted to another 

product like dimethyl ether, methyl formate, or acetic acid. 

Controlling and dissipating the heat of reaction and overcoming the equilibrium 

constraint to maximize the per-pass conversion efficiency are the two main process 

features that are considered when designing a methanol synthesis reactor, commonly 

referred to as a methanol converter. Numerous methanol converter designs have been 

commercialized over the years and these can be roughly separated into two categories - 

adiabatic or isothermal reactors. Adiabatic reactors often contain multiple catalyst beds 

separated by gas cooling devices, either direct heat exchange or injection of cooled, fresh 

or recycled, syngas. Axial temperature profiles often have a sawtooth pattern that is low 

at the point of heat removal and increases linearly between the heat exchange sections. 

The isothermal reactors are designed to continuously remove the heat of reaction so they 

operate essentially like a heat exchanger with an isothermal axial temperature profile. 

The following section describes some of the many methanol converter designs found in 

the literature. 

4.4 Reactors 

One of the more widely used commercial isothermal methanol converters is the Lurgi 

Methanol Converter (Haid and Koss 2001; Supp and Quinkler 1984). It is a shell and 

tube design similar to their Fischer Tropsch (FT) reactor. The tubes contain a proprietary 

Lurgi methanol catalyst (Cu/ZnO/Cr2O3 + promoters) and are surrounded by boiling 

water for reaction heat removal. These units operate at 50-100 bar and 230-265°C. 

Varying the pressure of the boiling water controls the reactor temperature. Byproduct 

steam is produced at 40-50 bar and can be used to run the compressor or to provide heat 

for the distillation process. 

The ICI Low pressure Quench Converter (Pinto and Rogerson 1977; Rogerson 1971; 

Rogerson 1984) is the most widely used adiabatic methanol converter. It is operated at 

50-100 bar and 270°C. The Cu/ZnO/Al2O3 catalyst is contained in a single bed supported 
33

by an inert material. Adding cold fresh and recycled syngas quenches the synthesis 

reaction and controls the reaction temperature. The gas is injected at appropriate depths 

within the reactor through spargers called lozenges. There are horizontal layers of these 

lozenges that run across the converter from side to side and each has an outer surface 

covered with wire mesh and a central pipe that delivers the cold gas. ICI has an 

improved version of this reactor known as an ARC converter. The main technical 

difference is that instead of a single continuous catalyst bed, the bed is separated by 

distribution plates to form multiple consecutive catalyst domains. 

Kellogg, Brown, and Root (now Halliburton) has developed an adiabatic methanol 

converter that has multiple fixed bed reactors arranged in series and separated by heat 

exchangers. All of the recycled syngas is fed directly into the first reactor stage. The 

reactors have a spherical geometry to reduce construction costs and they also use less 

catalyst compared to the ICI Quench Converter. The Haldor-Topsoe Collect, Mix, 

Distribute (CMD) converter operates on a similar principle. Vertical support beams 

separate catalyst beds. The gas inlet at the bottom of the reactor provides fresh syngas 

that flows radially up through the first catalyst bed. At the top of the reactor, this first 

pass through gas is mixed with quench gas and distributed evenly so that it flows radially 

down through the second catalyst bed (Dybkjaer 1981). The cited benefit of this design 

is an increase in per-pass conversion. Toyo Engineering Corporation has designed 

another version of a multistage radial flow methanol converter (MRF-ZTM) that uses 

bayonet boiler tubes for intermediate cooling. The tubes divide the catalyst into 

concentric beds. 

The Tube Cooled Converter is a reactor design that is simple to operate. The syngas 

enters at the bottom of the reactor where a manifold distributes the gas through tubes that 

act as a heat exchanger prior to the gas entering the catalyst bed. The Linde isothermal 

reactor, known as the Variobar converter, has the unique design feature of using coiled 

helical tubes embedded in the catalyst bed for heat removal. Spacers separate the multilayer coils and boiling water is circulated through the tubes. Mitsubishi Gas Chemical in 

collaboration with Mitsubishi Heavy Industry has developed an isothermal reactor known 

as the MGC/MHI Superconverter (Takase and Niwa 1985). This reactor design uses 

double-walled tubes that are filled with catalyst in the annular space between the inner 

and outer tubes. The feed syngas enters the inner tubes and is heated as it progresses 

through the tube. The gas then passes downward through the catalyst bed in the annular 

space. Heat is removed on both sides of the catalyst bed by the boiling water surrounding 

the tubes as well as by the feed gas introduced into the inner tube (Tijm et al. 2001). A 

high conversion rate (about 14 % methanol in the reactor outlet) is cited for this reactor 

(Fiedler et al. 2003). 

Additional methanol converter designs include technologies using three phase systems 

similar in principle to the slurry reactors used for Fischer Tropsch synthesis (FTS). 

These technologies are collectively known as Liquid Phase Methanol Synthesis. 

ChemSystems, Inc. and Air Products and Chemicals, Inc. with Department of Energy 

funding developed a liquid-entrained catalytic reactor for converting low H2/(CO + CO2) 

ratio syngas into methanol known as LPMEOHTM (Bhatt et al. 1999; DOE 1992). The 

ability to convert low stoichiometric ratio (CO rich) syngas lends itself to using syngas 
34

from coal or biomass gasification for methanol production. The three-phase slurry 

reactor provides better temperature control by uniformly dissipating the heat of reaction 

into the high heat capacity liquid. The LPMEOHTM process uses a supported Cu/ZnO 

catalyst (20-45 wt%) dispersed in circulating mineral oil with reactor temperatures of 

225-265°C and a pressure of 50 bar. Brookhaven National Laboratory has also 

developed a liquid phase methanol synthesis process that can be operated at lower 

pressures (< 5 atm) with up to 90% conversion without syngas recycling (Tijm et al. 

2001). This work has been done in collaboration with Amoco and in 1998 a 50 ml minipilot plant was successfully operated. 

Two other methanol conversion processes are based on systems in which the product 

methanol is continuously removed from the gas phase by selective adsorption on a solid 

or in a liquid. The Gas-Solid-Solid Trickle Flow Reactor (GSSTFR) utilizes an adsorbent 

such as SiO2/Al2O3 to trap the product methanol (Herman 1991; Pass et al. 1990). The 

solid adsorbent is collected in holding tanks and the methanol is desorbed. In the Reactor 

System with Interstage Product Removal (RSIPR), a liquid solvent is used to adsorb the 

product methanol (Herman 1991). 

4.5 Chemistry 

Catalytic methanol synthesis from syngas is a classic high-temperature, high-pressure 

exothermic equilibrium limited synthesis reaction. The chemistry of methanol synthesis 

is as follows (Rostrup-Nielsen 2000): 

CO + 2H2 Æ CH3OH ∆Hr = -90.64 kJ/mol 

CO2 + 3H2 Æ CH3OH + H2O ∆Hr = -49.67 kJ/mol 

CO + H2O Æ CO2 + H2 ∆Hr = -41.47 kJ/mol 

For methanol synthesis, a stoichiometric ratio, defined as (H2 – CO2)/(CO + CO2), of 

about 2 is preferred. This means that there will be just the stoichiometric amount of 

hydrogen needed for methanol synthesis. For kinetic reasons and in order to control byproducts, a value slightly above 2 is normally preferred (Dybkjaer and Christensen 2001). 

Although methanol is made from mixtures of H2 and CO, the reaction is about 100 times 

slower than when CO2 is present (Wender 1996). Until as recently as the 1990s the role 

of CO2 in methanol synthesis was not clear. The water gas shift (WGS) activity of Cu 

catalysts is so high that it was difficult to deconvolute the role of CO and CO2 in 

methanol synthesis. Isotopic labeling studies have unequivocally proved that CO2 is the 

source of C in methanol (Ladebeck 1993; Wender 1996). CO is involved in the reverse 

WGS reaction to make H2 and CO2. CO2 is also thought to keep the catalyst in an 

intermediate oxidation state Cu0

/Cu+ preventing ZnO reduction followed by brass 

formation (Ladebeck 1993). The proposed mechanism for catalytic methanol synthesis is 

believed to proceed through a long-lived formate intermediate. CO2 is adsorbed on a 

partially oxidized metal surface as a carbonate and hydrogenated. This intermediate is 
35

then hydrogenated in the rate-limiting step. The copper catalyst sites have high activity 

for splitting the first C-O bond in CO2 that helps maintain the oxidation state of the active 

copper sites. At high concentrations, however, CO2 actually reduces catalyst activity by 

inhibiting methanol synthesis. The feed gas composition for methanol synthesis is 

typically adjusted to contain 4-8% CO2 for maximum activity and selectivity. Even 

though Cu has WGS activity, excessive amounts of H2O also leads to active site blocking 

that is poor for activity but improves selectivity by reducing byproduct formation by 50% 

(Chinchen et al. 1990). 

4.6 Catalysts 

The first high-temperature, high-pressure methanol synthesis catalysts were ZnO/Cr2O3

and were operated at 350°C and 250-350 bar. Catalyst compositions contained 20-75 

atom% Zn. These catalysts demonstrated high activity and selectivity for methanol 

synthesis and proved robust enough to resist sulfur poisoning which is inherent when 

generating syngas from coal gasification. Over the years, as gas purification technologies 

improved (i.e., removal of impurities such as sulfur, chlorine, and metals), interest in the 

easily poisoned Cu catalysts for methanol synthesis was renewed. In 1966, ICI 

introduced a new, more active Cu/ZnO/Al2O3 catalyst that began a new generation of 

methanol production by using a low temperature (220-275°C), low pressure process (50-

100 bar). The last high temperature methanol synthesis plant closed in the mid-1980s 

(Fiedler et al. 2003) and, at present, low temperature low pressure processes based on Cu 

catalysts are used for all commercial production of methanol from syngas. The synthesis 

process has been optimized to point that modern methanol plants yield 1 kg of MeOH 

/liter of catalyst/hr with >99.5% selectivity for methanol. Commercial methanol 

synthesis catalysts have lifetimes on the order of 3-5 years under normal operating 

conditions. 

The Cu crystallites in methanol synthesis catalysts have been identified as the active 

catalytic sites although the actual state (oxide, metallic, etc.) of the active Cu site is still 

being debated. The most active catalysts all have high Cu content, optimum ~60% Cu 

that is limited by the need to have enough refractory oxide to prevent sintering of the Cu 

crystallites. The ZnO in the catalyst formulation creates a high Cu metal surface area; it 

is suitably refractory at methanol synthesis temperatures and hinders the agglomeration 

of Cu particles. ZnO also interacts with Al2O3 to form a spinel that provides a robust 

catalyst support. Acidic materials like alumina, are known to catalyze methanol 

dehydration reactions to produce DME. By interacting with the Al2O3 support material, 

the ZnO effectively improves methanol selectivity by reducing the potential for DME 

formation. 

Catalysts are typically prepared by the co-precipitation of metal salts (such as nitrates or 

sulfates) with a variety of precipitating agents. It is important to avoid contaminating 

methanol catalysts with metals that have FT activity (Fe or Ni) during synthesis. 

Incorporation of alkali metals in catalyst formulations should also be avoided because 

they catalyze the increased production of higher alcohols. Table 11 shows catalyst 

formulations for the many commercial manufacturers. 
36

Table 11: Commercial Methanol Synthesis Catalyst Formulations 

Manufacturer Cu at% Zn at% Al at% Other Patent date 

IFP 45-70 15-35 4-20 Zr – 2-18 1987 

ICI 20-35 15-50 4-20 Mg 1965 

BASF 38.5 48.8 12.9 1978 

Shell 71 24 Rare Earth 

oxide –5 

1973 

Sud Chemie 65 22 12 1987 

Dupont 50 19 31 none found 

United Catalysts 62 21 17 none found 

Haldor Topsoe 

MK-121 

> 55 21-25 8-10 none found 

Additional catalyst formulations have been presented in the literature with the purpose of 

improving per-pass methanol yields (Klier 1982). The addition of Cs to Cu/ZnO 

mixtures has shown improved methanol synthesis yields. This only holds true for the 

heavier alkali metals, as the addition of K to methanol synthesis catalysts tends to 

enhance higher alcohol yields. Cu/ThO2 intermetallic catalysts have also been 

investigated for methanol synthesis (Klier 1982). These catalysts have demonstrated high 

activity for forming methanol from CO2-free syngas. Thoria based methanol catalysts 

deactivate very rapidly in the presence of CO2. Cu/Zr catalysts have proven active for 

methanol synthesis in CO-free syngas at 5 atm and 160-300°C (Herman 1991). 

Supported Pd catalysts have also demonstrated methanol synthesis activity in CO2-free 

syngas a 5 -110 atm at 260-350°C. 

4.7 Gas Cleanliness Requirements 

Through the years, methanol catalysts have been optimized for maximum selectivity, 

spacetime yields, and long service life, but they still deactivate (Wender 1996). 

Nevertheless, under normal commercial operating conditions, CuO/ZnO methanol 

catalysts have quite long lifetimes, up to 2-5 years. Increasing the process operating 

temperature to maintain activity and yield compensates for progressive catalyst 

deactivation. 

Copper has a very low activity for breaking C-O bonds or forming C-C bonds. As a 

result, in methanol synthesis, wax formation is not usually a major problem nor is 

formation of coke from hydrocarbons (Twigg and Spencer 2001). Cu catalysts also do 

not have a strong tendency to catalyze FT reactions, or processes involving carbonium 

ion chemistry. Therefore, the major causes of methanol catalyst deactivation are 

poisoning and Cu sintering that results in loss of Cu surface area. 

Copper catalysts are extremely sensitive to site-blocking poisons such as reduced sulfur. 

To retain the long-term activity of Cu catalysts, it has been found empirically that the 

gas-phase sulfur concentration needs to be kept below 1 ppm and preferably below 0.1 
37

ppm. High partial pressures of water vapor also inhibit the reaction, presumably by 

competitive adsorption on active sites. 

Interestingly enough, the catalyst’s formulation itself tends to reduce the effects of sulfur 

poisoning. The ZnO in the catalyst is very effective in limiting poisoning of copper 

catalysts because it acts as a poison sink, diverting the poisons away from the active Cu 

sites and scavenging the sulfur out of the gas stream by forming ZnS and ZnSO4. 

Commercial Cu-ZnO catalysts can absorb 0.4 wt% sulfur and still maintain 70% of their 

activity (Kung 1992). This helps to extend the catalyst lifetime. 

As it turns out, COS is not a gas phase methanol catalyst poison. Studies have shown 

that COS does not cause any deactivation over a concentration range of 0.6-9 ppm. The 

Cu/ZnO catalysts have no activity for COS adsorption or dissociation and at methanol 

synthesis conditions, COS does not react with H2 to form H2S (Kung 1992). Gas phase 

COS does not pose a catalyst poisoning problem, however, it becomes a serious catalyst 

poison in the liquid phase methanol synthesis process.f 

If reactor temperatures are not properly controlled, the highly exothermic methanol 

synthesis reactions can rapidly cause catalyst deactivation by sintering of the Cu 

crystallites. Reactor temperatures are maintained below 300°C to minimize sintering. 

The presence of Cl in syngas has also been correlated with a greatly enhanced rate of 

sintering of copper crystallites most likely due to the formation of volatile copper 

chloride compounds. The mechanism of Cl poisoning can be explained by adsorbed Cl 

atoms that block or modify active catalyst sites. The Cl interacts with the Cu by forming 

CuCl2 that has a low melting point and high surface mobility that accelerates catalysts 

sintering. The presence of trace amounts of chlorine also causes enhanced sulfur 

poisoning, probably because Zn chlorides also form removing these sites as possible 

sinks for sulfur capture. The limits on HCl content to avoid catalyst poisoning are more 

severe than for H2S poisoning, on the order of 1 ppb (Twigg and Spencer 2001). 

Other gas phase poisons in syngas that need to be avoided are metal carbonyls, 

particularly Ni and Fe carbonyls. Trace amounts of metal carbonyl compounds could be 

present in feed streams, especially in laboratory experiments where stainless steel tubing 

is used. Volatile metal carbonyl compounds deposit on the catalyst surface and block 

active sites. The deposition of Fe and Ni carbonyls affects the selectivity of the methanol 

catalysts. Ni increases the production of CH4 and Fe increases the production of FT 

products. Metal carbonyl concentrations should be below 5 ppb (Wender 1996). 

The purity of catalyst formulations is also critical to insure optimum activity and 

selectivity of methanol synthesis catalysts. It is very important to eliminate Ni and Fe 

impurities in catalyst precursors during catalyst formulation. These impurities in 

methanol synthesis catalysts reduce the selectivity and increase the yield of CH4 and FT 

hydrocarbon products. Alkali metals should also be avoided because these reduce overall 

catalyst activity and increase the production of higher alcohols. Unwanted SiO2 in the 

catalyst promotes wax formation and unreacted (non-spinel) alumina enhances DME 

formation. Other poisons to be avoided are As and P. 
38

Maintaining low levels of sulfur in the syngas used to produce methanol is typically not a 

problem. Most syngas used for methanol synthesis is produced from steam reforming of 

methane over Ni catalysts that are very intolerant of sulfur. Therefore, sulfur compounds in 

the feedstock will normally be reduced to below 0.1 ppm by desulfurization prior to the 

primary reformer in hydrogen, ammonia, and methanol plants. Poisoning by chlorine or 

other compounds is also negligible in normal plant operation, because these are removed 

by the earlier stages of generation of synthesis gas. A summary of the gas cleanliness 

requirements for gas phase and liquid phase methanol production are in Table 12. 

Table 12: MeOH Gaseous Contaminant Constraints 

Gas phase ppmv Reference LPMeOH (Novem 2000) ppmv 

Sulfur 

(not COS) 

<0.5 (<0.1 

preferred) 

Kung, 1992 Sulfur (including OCS) 0.1 

Halides 0.001 Twigg and 

Spencer 2001 

Total halides 0.01 

Fe and Ni 0.005 Kung, 1992 Acetylene 5 

Total unsaturates 300 

NH3 10 

HCN 0.01 

Fe and Ni 0.01 

4.8 Status/Technology Developers 

BASF of Germany introduced the first commercial methanol synthesis route in 1923 at 

pressures of 250-350 atm (3,675-5,145 psi). Plants operated at high pressures until lowpressure routes were developed and by the early 1980s the majority of the producers had 

switched from the high-pressure process to a low pressure one. This was done because 

the low-pressure process is more efficient, has a lower capital cost, and is less expensive 

to operate. ICI and Lurgi offer the two most widely used low-pressure processes 

(Davenport 2002). Low-pressure technologies are also offered by Mitsubishi Gas 

Chemical Company, Inc. of Japan, jointly by Haldor-Topsoe of Denmark and Nihon 

Suiso Kogyo of Japan, and by Halliburton. Table 13 presents the process conditions for 

each supplier. 
39

Table 13: Methanol Technology Suppliers 

Technology 

supplier 

T (C) P (atm) Notes 

ICI (Synetix) 210-290 50-100 Currently licenses 4 types of reactors: ARC, Tubular 

Cooled, Isothermal Linde, and Toyo (see reactor 

section for details) 

Lurgi 230-265 50-100 Tubular, isothermal reactor 

Mitsubishi 240 77-97 Tubular, isothermal reactor 

Haldor-Topsoe & 

Nihon 

260 48-300 To date, no commercial plants based on this process. 

Halliburton 

(Kellogg, Brown & 

Root) 

not found not found Spherical reactor geometry 

Source: (Davenport 2002; Dybkjaer and Christensen 2001; Fiedler et al. 2003; Lee 1990; Takase and Niwa 

1985; Tijm et al. 2001) 

The world’s largest producer of methanol is Methanex; they account for 17% of the total 

global capacity (Davenport 2002). The next largest producer is SABIC which accounts 

for 6.5% of the total global capacity (Davenport 2002). As of January 1, 2002, 

worldwide annual capacity of methanol was 12.8 billion gallons (38 million tonnes, 9 

million GJ) (Davenport 2002). In 2001, the U.S. consumed 28% of the total world 

methanol while North America produced 20% of the world supply (Davenport 2002). 

There are 18 methanol production plants in the U.S. (Lidderdale 2001). 

Recently, Lurgi won a contract to build a methanol plant for National Petrochemical's 

subsidiary Zagros Petrochemical in Bandar Assaluye, Iran. The capacity will be 1.8 

million tonnes/yr., making it the largest methanol unit in the world. It is expected to 

begin operation in 2004 [http://www.qipc.net/news/june/jun0051.htm]. 

Liquid phase methanol synthesis (LPMeOH) was invented in late 1970s (Hamelinck and 

Faaij 2001) and developed in the 1980s (Bhatt et al. 1999). In the 1980s, Air Products 

and Chemicals, Inc. tested this technology at a pilot-scale facility with an output of 8 tpd 

of methanol in LaPorte, TX. Later, a 260 TDP LPMeOH demonstration plant was built 

in Kingsport, TN and began operation in 1997 (Bhatt et al. 1999). It was a joint venture 

between Eastman Chemical Company and Air Products and Chemicals, Inc. (Wender 

1996). 

4.9 Process Performance 

Converting syngas to methanol in commercial units is limited by thermodynamic 

considerations to about 25% per-pass (Wender 1996). Lurgi’s multitubular reactor 

produces close to 1 kg of methanol per liter of catalyst per hour. The actual per-pass 

conversion of methanol is only 4-7 vol% (Wender 1996). Removal of methanol would 

increase per-pass conversions. Another way to increase the conversion is in situ 

dehydration of methanol to DME. Generally, 100 tons of methanol is converted to nearly 

44 tons of hydrocarbons and 56 tons of water (Maiden 1983; Wender 1996). 
40

4.10 Uses 

Globally, formaldehyde production is the largest consumer of methanol (35% of 

worldwide methanol) (Davenport 2002). This is followed by MTBE at 25%. The third 

largest consumer is acetic acid at 9% (Davenport 2002). The following sections discuss 

the uses of methanol. 

4.10.1 Formaldehyde 

Formaldehyde was first commercially produced in 1890 in Germany by the catalytic 

partial oxidation of methanol with air over an unsupported Cu catalyst at atmospheric 

pressure (Satterfield 1991). By the beginning of the 20th century, Cu was replaced by Ag 

as the catalyst of choice because yields were better. Use of Iron molybdate catalysts 

began in the 1940s (Satterfield 1991). 

Formaldehyde produced over Ag catalysts proceeds via methanol dehydrogenation and 

partial oxidation according to the following reactions: (Chen and Masel 1995; Ruf and 

Emig 1999): 

CH3OH Æ H2CO + H2 (dehydrogenation) ∆Hr = +84 kJ/mol 

CH3OH + 1⁄2 O2 Æ H2CO + H2O (partial oxidation) ∆Hr = -159 kJ/mol 

For the Ag catalyst process, a 50% methanol mixture in air is passed through a thin, fixed 

bed at 600°C and slightly above atmospheric pressure. Metallic Ag is not active for 

methanol decomposition. The active Ag site requires chemisorbed oxygen, which 

provides a site for methanol adsorption. 

For the iron molybdate [Fe2(MoO4)3] process, a lean methanol-air mixture is used. This 

reaction is more exothermic than the Ag process so heat removal is required (Satterfield 

1991). The catalyst temperature is about 340-380°C. Excessive reactor temperatures 

cause volatilization of molybdenum oxide, which reduces the selectivity of the process. 

A high oxygen partial pressure is required to maintain catalyst activity and in the 

presence of excess methanol catalyst activity is lost. 

Formaldehyde is produced commercially from methanol by three industrial processes 

(Reuss et al. 2003): 

1. Partial oxidation and dehydrogenation with air in the presence of silver crystals, steam, 

and excess methanol at a temperature of 680-720 °C, otherwise known as the BASF 

process. The methanol conversion for this process is 97-98%. 

2. The same as process 1 except in the presence of crystalline silver or silver gauze at a 

temperature of 600-650 °C. Then, the product is distilled and the unreacted methanol is 

recycled. The primary methanol conversion for this process is 77-87%. 
41

3. Oxidation with only excess air in the presence of a modified 

iron/molybdenum/vanadium oxide catalyst at a temperature of 250-400 °C, also know as 

the Formox process. The methanol conversion for this process is 98-99%. 

Formaldehyde is used to make resins with phenol, urea, or melamine for the manufacture 

of various construction board products. The demand for formaldehyde is driven by the 

construction industry. 

4.10.2 MTBE (Methyl Tertiary-Butyl Ether) 

In general, greater than 95% of the MTBE produced is used in the gasoline pool. MTBE 

is also used in the petrochemical industry for the production of isobutene. Additionally, 

it can be used in a number of chemical reactions including methacrolein, methycrylic 

acid, or isoprene production. One other use is solvent dewaxing of hydrocarbon oils 

(Peters et al. 2003). 

MTBE is produced by reacting isobutene with methanol in the presence of an acidic 

catalyst according to the following slightly exothermic reaction: 

i-C4H8 + CH3OH Æ (CH3)3COCH3 ∆Hr = -37 kJ/mol 

The reaction temperature and pressure are 30-100°C and 7-14 atm (100-200 psig), 

respectively (Schädlich et al. 2003), so that the reaction occurs in the liquid phase. 

Catalysts used are solid acids, zeolites (H-ZSM-5), and macroporous sulfonic acid ionexchange resins such as Amberlyst-15 (Collignon et al. 1999). A molar excess of 

methanol is used to increase isobutene conversion and inhibit the dimerization and 

oligomerization of isobutene. At optimum reaction conditions, MTBE yields approaching 

90% can be achieved. 

In general, most commercially available processes are similar, consisting of a reaction 

and a refining section. In addition to the Snamprogetti and Hüls (now Oxeno) processes, 

processes have been developed by Arco, IFP, CDTECH (ABB Lummus Crest and 

Chemical Research Licensing), DEA (formerly Deutsche Texaco), Shell (Netherlands), 

Phillips Petroleum, and Sumitomo (Peters et al. 2003). Currently, there are over 140 

MTBE plants with a total installed capacity of about 20 million tonnes/yr. 

The largest demand for MTBE has [primarily] been in the U.S., with North America 

consuming roughly 65% of the worldwide production of MTBE. U.S. consumption 

increased substantially with the Clean Air Act Amendments of 1990. Recently, however, 

MTBE has come under environmental attack because of leaky tanks and groundwater 

contamination. In fact, California planned to ban MTBE from gasoline by January 1, 

2003 but the ban has been pushed back a year. Recently, Alberta Envirofuels closed its 

800,000 tonne/yr MTBE plant located in Edmonton, Canada (2002). They plan to 

convert the plant into an iso-octane facility (a gasoline additive that is an alternative to 

MTBE) at a cost of $50 million. With the reduction and possibly elimination of MTBE 

in gasoline, a worldwide decline in the demand for MTBE of about 3-5% is predicted 

(Davenport 2002). 
42

4.10.3 Acetic Acid 

The third most abundant chemical synthesized from methanol is acetic acid (CH3OH + 

CO Æ CH3COOH). Vinyl acetate, acetic anhydride and terephthalic acid are all 

manufactured from acetic acid. Latex emulsion resins for paints, adhesives, paper 

coatings, and textile finishing agents are made from vinyl acetate. Acetic anhydride is 

used in making cellulose acetate fibers and cellulosic plastics. In terms of world 

production of acetic acid, approximately one-half comes from methanol carbonylation 

and about one-third from acetaldehyde oxidation (Wagner 2002). Cheung et al. 2003 

states that there is little doubt that the technology of the future is methanol carbonylation. 

In 1999, the worldwide demand for acetic acid was 2.8 million tonnes (6.17 × 109 lb) 

(Wagner 2002). 

Acetic acid is produced by the carbonylation of methanol with CO in the liquid phase 

using Co, Rh, or Ni catalysts promoted by iodine. The synthesis of acetic acid by 

carbonylation of methanol is one of the most important industrial applications of 

homogeneous catalysis (Roper 1991). Two processes have been developed based on this 

homogeneous catalytic synthesis. One is the BASF process (which is the older of the 

two) which uses a Co/iodine catalyst at process conditions of 250°C and 500-700 atm 

(3,000-10,000 psig) with 90% selectivity to acetic acid (from methanol). The second is 

the Monsanto process, which uses an Rh/iodine catalyst at process conditions of 180°C 

and 30-40 atm with over 99% selectivity. This liquid phase process is initiated by the 

reaction of methanol with HI to yield methyl iodide. The active catalyst in the process is 

the metal carbonyl [RhI2(CO)2]

-

 into which methyl iodide is inserted in the rate-limiting 

step. Acetic acid is formed by the hydrolysis of the eliminated acetyl iodide species 

CH3COI that also regenerates HI (King and Grate 1985). 

The second, less severe Monsanto process has displaced the BASF process and all new 

installed capacity since 1973 is based on the Monsanto process (SRI 1994). The catalytic 

system is very corrosive and requires expensive steels for materials of construction. 

Complete recovery of the expensive Rh catalyst (~10-3 M) and recycle of HI (~0.1 M) is 

paramount to maintain favorable process economics. The high cost of Rh has lead to the 

search for other, lower cost, metals that could be used as acetic acid catalysts with similar 

performance. 

4.10.4 Methanol to Gasoline (MTG) 

The methanol to gasoline (MTG) process developed by Mobil Oil Corporation involves 

the conversion of methanol to hydrocarbons over zeolite catalysts. The MTG process, 

although considered the first major new synthetic fuel development since FTS, was 

discovered by accident in the 1970s by two independent groups of Mobil scientists trying 

to convert methanol to ethylene oxide and attempting to methylate isobutene with 

methanol over a ZSM-5 zeolite catalyst (Keil 1999). 

The MTG process occurs in two steps. First, crude methanol (17% water) is super-heated 

to 300°C and partially dehydrated over an alumina catalyst at 27 atm to yield an 

equilibrium mixture of methanol, dimethyl ether, and water (75% of the methanol is 
43

converted). This effluent is then mixed with heated recycled syngas and introduced into 

a reactor containing ZSM-5 zeolite catalyst at 350-366°C and 19-23 atm to produce 

hydrocarbons (44%) and water (56%) (Hancock 1985). The overall MTG process 

usually contains multiple gasoline conversion reactors in parallel because the zeolites 

have to be regenerated frequently to burn off the coke formed during the reaction. The 

reactors are then cycled so that individual reactors can be regenerated without stopping 

the process, usually every 2-6 weeks (Kam et al. 1984). 

The MTG reactions may be summarized as follows (Wender 1996): 

2CH3OH Æ CH3OCH3 + H2O 

CH3OCH3 Æ C2-C5 olefins 

C2-C5 olefins Æ paraffins, cycloparaffins, aromatics 

The selectivity to gasoline range hydrocarbons is greater than 85% with the remainder of 

the product being primarily LPG (Wender 1996). Nearly 40% of the gasoline produced 

from the MTG process is aromatic hydrocarbons with the following distribution: 4% 

benzene, 26% toluene, 2% ethylbenzene, 43% xylenes, 14% trimethyl substituted 

benzenes, plus 12% other aromatics (Wender 1996). The shape selectivity of the zeolite 

catalyst results in a relatively high durene (1,2,4,5-tetramethylbenzene) concentration, 3-

5% of the gasoline produced (MacDougall 1991). Therefore, MTG gasoline is usually 

distilled and the heavy fraction is processed in the heavy gasoline treating unit to reduce 

the durene concentration to below 2%. This results in a high quality gasoline with a high 

octane number. The MTG process may no longer be a good option for gasoline 

production because the 1990 Clean Air Act Amendments limits the amount of aromatics 

in reformulated gasoline in order to reduce air toxics. In California, the amount of total 

aromatics is limited to 25 vol% (Owen and Coley 1995). 

The first commercial MTG plant came onstream in 1985 in New Zealand (Mobil’s 

Motunui plant), producing both methanol and high octane gasoline from natural gas. 

Plant production averaged 14,500 BPD of gasoline until 1997, at which time gasoline 

manufacturing was abandoned and the plant produced only methanol. No information 

could be found regarding the reason for abandoning the manufacture of gasoline, but it is 

most likely due to economic factors such as the cost of oil and the market price of 

methanol. Presently this plant, along with a nearby methanol plant at Waitara produces, 

2.43 million tonnes per year of chemical grade methanol for export 

[http://www.techhistory.co.nz/pages/Petrochemical%20Decisions.htm]. 

A fluid bed MTG plant was jointly designed and operated near Cologne, Germany by 

Mobil Research and Development Corp., Union Rheinische Braunkohlen Kraftstoff AG 

and Uhde Gmb (Keil 1999). A demonstration plant (15.9 m3

/day) operated from 1982 to 

1985. Although, no commercial plants have been built, the fluid bed technology is ready 

for commercialization. 
44

4.10.5 Methanol to Olefins (MTO) and Methanol to Gasoline and Diesel (MOGD) 

Along with the MTG process, Mobil developed several other processes for converting 

methanol to hydrocarbons based on zeolite catalysts. Since light olefins are intermediates 

in the MTG process, it is possible to optimize the methanol to olefins (MTO) synthesis. 

Higher reaction temperatures (~500°C), lower pressures, and lower catalyst (acidity) 

activity favors light olefin production (Keil 1999). The rate of olefin production could be 

modified so that 80% of the product consists of C2 to C5 olefins rich in propylene (32%) 

and butenes (20%) with an aromatic rich C5+ gasoline fraction (36%) (MacDougall 1991; 

Wender 1996). The process can also be modified for high ethylene and propylene yield 

(>60%). 

Mobil also developed a methanol to gasoline and diesel (MOGD) process. 

Oligomerization, disproportionation and aromatization of the olefins produced in the 

MTO synthesis are the basis for the MOGD process. In the MOGD process, the 

selectivity of gasoline and distillate from olefins is greater than 95% (Keil 1999). One 

source gives the gasoline product from MOGD to be: 3 wt% paraffins, 94 wt% olefins, 1 

wt% napthenes, and 2 wt% aromatics (Tabak and Yurchak 1990). A large-scale test run 

was performed at a Mobil refinery in 1981. 

Neither the MTO nor the MOGD process is currently in commercial practice (Wender 

1996). However, UOP and HYDRO of Norway, does license their own methanol to 

olefins process where the primary products are ethylene and propylene 

[http://www.uop.com/petrochemicals/processes_products/mto_intro.htm and (Keil 

1999)]. They use a fluidized bed reactor at 400 – 450°C and achieve roughly 80% carbon 

selectivity to olefins at nearly complete methanol conversion (Apanel and Netzer 2002). 

The operating parameters can be adjusted so that either more ethylene is produced (48 

wt% ethylene, 31 wt% propylene, 9 wt% butenes and 1.5 wt% other olefins) or else more 

propylene (45 wt% propylene, 34 wt% ethylene, 12 wt% butenes and 0.75 wt% other 

olefins). 

4.10.6 TIGAS (Topsoe Integrated Gasoline Synthesis) 

The Topsoe Integrated Gasoline Synthesis (TIGAS) process was developed by Haldor 

Topsoe with the intent of minimizing capital and energy costs by integrating methanol 

synthesis with the MTG step into a single loop, without isolation of methanol as an 

intermediate (Keil 1999; Topp-Jorgensen 1987; Topp-Jorgensen 1988). This process was 

developed with the idea that future plants would be constructed in remote areas for 

recovery of low cost natural gas. In Mobil’s MTG process, different pressures are 

preferred for syngas production, methanol synthesis, and the fixed bed MTG step. These 

pressures are 15-20 atm (221-294 psi), 50-100 atm (735-1470 psi), and 15-25 atm (221-

368 psi), respectively (Wender 1996). The TIGAS process involves modified catalysts 

and conditions so that the system pressure levels out and separate compression steps are 

not required. In order to do this, a mixture of methanol and DME is made prior to 

gasoline synthesis. This approach results in only one recycle loop which goes from the 

gasoline synthesis step back to the MeOH/DME synthesis step. A 1 MTPD 

demonstration plant was built in Houston, Texas in 1984 and operated for 3 years (Topp-
45

Jorgensen 1988). The gasoline yield for the TIGAS process, defined as the amount of 

gasoline produced divided by the amount of natural gas feed and fuel, was shown to be 

56.5 wt% (Topp-Jorgensen 1988). Because of a reduced selectivity to gasoline range 

aromatics, the TIGAS process, however, yields a lower quality gasoline with a lower 

octane number, compared to MTG (MacDougall 1991). 

4.10.7 DME (Dimethyl Ether) 

DME is industrially important as the starting material in the production of the 

methylating agent dimethyl sulfate. It is also used as an aerosol propellant. DME has the 

potential to be used as a diesel or cooking fuel, a refrigerant, or a chemical feedstock 

(Gunda et al. 1995; Peng et al. 1999b; Shikada et al. 1999). Commercial production of 

DME originated as a byproduct of high-pressure methanol production. 

DME is formed in a two-step process where first, methanol is synthesized, then it is 

dehydrated over an acid catalyst such as γ-alumina at methanol synthesis conditions. The 

DME reaction scheme is as follows (Hansen and Joensen 1991; Peng et al. 1999a): 

CO + 2H2 Æ CH3OH (methanol synthesis) ∆H = -21.6kcal/mole 

2CH3OH Æ CH3OCH3 + H2O (methanol dehydration) ∆H = -5.6 kcal/mole 

H2O + CO Æ H2 + CO2 (water gas shift) ∆H = -9.8 kcal/mole 

NET reaction: 3H2 + 3CO Æ CH3OCH3 +CO2 ∆H = -58.6 kcal/mole 

Note that one product in each reaction is consumed by another reaction. Because of the 

synergy between these reactions, syngas conversion to DME gives higher conversions 

than syngas conversion to methanol. Shikada et al. 1999 gives the per-pass and total 

conversion for the synthesis of methanol, methanol/DME and DME to be: 

Table 14: Conversions for Methanol, Methanol/DME, and DME 

Conversion MeOH MeOH/DME DME 

Per-pass (%) 14 18 50 

Total (%) 77 85 95 

The optimum H2:CO ratio for DME synthesis is lower than that for methanol synthesis 

and,ideally, should be around one (Peng et al. 1999a; Peng et al. 1999b; Shikada et al. 

1999). Recent improvements to the DME synthesis process involve the development of 

bifunctional catalysts to produce DME in a single gas phase step (i.e., one reactor) (Ge et 

al. 1998; Peng et al. 1999b) or the use of a slurry reactor for liquid phase dimethyl ether 

(LPDME) synthesis (Brown et al. 1991; Sardesai and Lee 1998). 
46

4.10.8 M100 and M85 

Neat methanol (M100) has many desirable properties that make it an attractive alternative 

to petroleum gasoline. It has a high heat of vaporization and relatively low heating value 

(about half of gasoline) that leads to a lower flame temperature compared to gasoline. 

This results in lower CO, NOx, and hydrocarbons emissions, but higher formaldehyde 

emissions. In general, neat methanol has only found use as a motor fuel in highperformance racing engines and in airplanes that have been fully modified and adapted to 

operate on methanol (Fiedler et al. 2003). This includes replacing plastic components in 

the fuel system, modifying the carburetor or fuel injection system, and preheating the fuel 

mixture. 

Methanol is miscible in water and in low concentration methanol mixtures, phase 

separation can occur. Methanol is also more corrosive than gasoline, however, it is less 

toxic and is not carcinogenic. Methanol was used as a gasoline blending agent prior to 

the mid-1980s. After that time, the EPA prohibited use of methanol in unleaded gasoline 

without the use of a cosolvent alcohol (Davenport et al. 2002). This was done, primarily, 

because methanol exhibits poor solubility in gasoline and has phase separation problems. 

However, there has been some success with M85, a mixture of methanol and gasoline (85 

vol%/15 vol%, respectively). Methanol is most commonly used in a mixture with 15% 

gasoline. This is done for two reasons: (1) a methanol flame is colorless, so gasoline is 

added to give the flame some color so that a fire can be detected in the event of a vehicle 

crash and (2) neat methanol can cause cold-start problems in the winter, or vapor lock in 

the summer [http://www.altfuels.org/m85.html]. Although automobiles have been 

manufactured to operate on M85, and some fleets continue to use M85, this fuel has 

failed to develop on a large-scale. Contributing factors include adverse consumer 

perceptions coupled with a lack of sufficient infrastructure, even though only small 

changes are required at gasoline fueling station (e.g., linings and seals in tanks, pumps, 

and dispensers) in order to handle M85 (Davenport et al. 2002). 

M100 is also being targeted for use in direct methanol fuel cells (DMFC), for fuel cell 

vehicle applications. DMFCs allow the use of methanol as fuel without requiring a fuel 

processor to extract hydrogen from the methanol. Significant progress has been made on 

finding better electrolyte material to prevent methanol crossover through the membrane. 

Operating temperatures of direct methanol fuel cells has increased to nearly 100°C. 

Improved anode catalysts have facilitated methanol oxidation and eliminated the need for 

an on-board hydrogen reformer. Several automakers have developed methanol powered 

fuel cell vehicles and currently, DaimlerChrysler and Honda both have two prototype 

direct methanol fuel cell vehicles. In fact, in 2002 one of DaimlerChrysler’s vehicles 

(NECAR 5) completed a 3,000 miles cross-country trip across the U.S. 
47

4.11 Environmental Performance 

Methanol is stable under normal storage conditions. It is not subject to hazardous 

polymerization reactions, but can react violently with strong oxidizing agents. The 

greatest hazard involved in handling methanol is the danger of fire or explosion (English 

et al. 1995). 

When used in spark engines, the higher the concentration of methanol in a 

methanol/gasoline blend, the lower the amount of CO, HC, and NOx in the exhaust 

emissions (Letcher 1983). In fact, the addition of tertiary butyl alcohol (TBA) and 

methanol may reduce the CO emissions up to 40 % and the hydrocarbon emissions up to 

20 % in comparison with conventional gasoline (Schädlich et al. 2003). 

As far as emissions are concerned, the MTG product has nearly the same emissions as 

gasoline from oil unless the methanol for the process is formed from biomass (Vermillion 

et al. 2001). 

4.12 Cost of Methanol 

From January 1993 to October 2001, U.S. spot prices for methanol ranged from a low of 

about $0.20/gallon to a high of about $1.50/gallon (Davenport 2002). In general, the 

price tended to be between $0.30/gallon ($5/GJ, LHV) and $0.70/gallon ($12/GJ LHV). 

Price fluctuations are often a result of oversupply, but the price is also highly dependent 

on the availability of imports and the price of natural gas. The current price of methanol 

listed in Chemical Marketing Reporter (2003) for January 28th is $0.40/gallon ($7/GJ, 

LHV) given as a contract price at the point of production. However, on January 6th, the 

price was listed as $0.62/gallon ($10/GJ, LHV) (2003). 

In producing methanol from natural gas, the majority of the capital cost comes from the 

reforming and gas-conditioning steps. A capital cost breakdown for the ICI process is 

given as: desulfurization (2%), reforming/gas cooling (32%), steam production (14%), 

compression (24%), methanol synthesis (22%), and distillation (6%) (Wender 1996). 

Using a slurry reactor instead of tubular one reduces the capital cost and decreases the 

compression cost (i.e. lower pressure drop across the reactor). In fact, LPMEOH 

investment costs are expected to be 5-23% less than the gas phase process of the same 

capacity (Hamelinck and Faaij 2001). 

Vermillion et al. 2001, claims that if methanol is formed from biomass, the cost would be 

$0.30/gallon to $0.55/gallon higher than if it were produced from natural gas. Table 15 

gives cost information from several coal based methanol studies and Table 16 gives the 

cost for several biomass-to-methanol studies. 
48

Table 15: Comparison of Coal Based Methanol Economic Studies 

Synthetic Fuels Associates, Inc. (1983) 

Study (ChemSystems 

1990) 

(Salmon 1986) 

max MeOH once through 

(Wham and Forrester III 1980) 

(Michae1979) 

Cost year 1987 1984 1982 1982 1979 1979 Coal feed rate (tonne/day) 

unclear 22,000-30,000 

(as received) 

23,054 (as received) 

9,662 (as received) 

18,331 (MAF) 

4,100 (as receiv

Coal cost $1.48/GJ 

$38.6/tonne 

$0.95/GJ and $1.42/GJ 

$46.47/tonne $50.99/tonne $0.95/GJ $16.5/ton

(mine mo

Methanol produced (tonne/day) 

886 plus 833 MWe 

5,000-10,000 14,439 2,231 

plus 870 MWe 

5,067 1,995 

Price of methanol (HHV basis) 

$6.6/GJ $0.45/gallon electricity sold at $0.37/kWh 

$7.2-$8.8/GJ $0.49-$0.60/gal 

$6.16/GJ $0.42/gallon (price was set and IRR calculated to be marginal at 7%) 

$5.87-$9.43/GJ $0.40-$0.64/gal electricity sold at $0.42/kWh 

$8.2/GJ $0.56/gallon 

$5/GJ $0.34/gal and $8/GJ $0.53/gal 

Level of detail Some details – 

Gives total capital investment and operating costs, and other economic assumptions 

Somewhat detailed – puts previous studies on same cost basis 

Some details – Gives total capital investment and operating costs, and other economic assumptions 

Somewhat detailed – Gives total capital investment, operating costs, and other economic assumptions 

Some detGives totainvestmeoperating and other economic assumptio
49

Table 16: Comparison of Biomass Based Methanol Studies 

Study (Hamelinck and 

Faaij 2001) (a), (b)

(Novem 2000) (a), (c)

(Larson 1992) (a) (Wyman et al. 

1993) (a)

(1990) 

Cost year 2001 1999 1991 1989 1987 Biomass feed rate (BD tonne/day) 

1,920 1,358 1,650 1,814 1,814 - 9,841 

Biomass cost $2/GJ 

$38/dry tonne 

$3/GJ $55/dry tonne 

$2/GJ $38/dry tonne 

$2.45/GJ $46/dry tonne 

$2.45/GJ $46/dtonne ($0-2.4/Gsensitivity) 

Electricity selling price 

$0.03/kWh $0.067/kWh $0.05/kWh 

(purchased) 

not given $0.041/kWh 

Net power (MW) no excess to about 100 

MW (@ 91% production cap) 

about 150 no excess no excess no excess 

Other fuels none co-fires nat gas in 

gas turbine 

none none none 

MeOH produced (tonne/day) 

613-970 (@ 91% production cap) 

57-196 705-1,004 790-1,110 717 – 4,577 

Price of MeOH (HHV basis) 

$9-$12/GJ $0.58-$0.83/gallon 

$10.6-$13.6/GJ $0.72-$0.93/gallon 

$9.62-$15.4/GJ $0.66-$1.05/gallon 

$10.01-$19.60/GJ $0.68-$1.34/gallon (@ 12% discount rate) 

$10-$18.9/GJ $0.68-$1.29/ga(for 20% capitarecovery factor

Level of detail Very detailed - Gives 

costs of individual equipment, operating costs, and other economic parameters and assumptions 

Some details - Gives costs of major plant sections, and assumptions used to determine installed costs 

Somewhat detailed - Gives capital and operating costs 

Somewhat detailed - Gives capital and operating costs and cash flow sheet 

Somewhat detailed - Gives capital cost by section and operating costs 

Notes: (a) Examined direct & indirect gasifiers and atmospheric and pressurized gasifiers. 

(b) Several configurations examined (varied type of gas cleaning, reforming, and gas or LPMeOH with and withou(c) Examined once through LPMeOH with combined heat and power. 
50

4.12.1 Cost of Formaldehyde 

A study by others was performed to examine the three industrial routes for producing 

formaldehyde. The cost of production for each process is presented in Table 17 (Reuss et 

al. 2003). The BASF process is the most economical, followed by the Formox process. 

The key feature in the BASF process design is a liquid recirculation system that produces 

excess steam while simultaneous savings cooling water. For comparison, the price of 

formaldehyde given in Chemical Marketing Reporter for January 28th 2003 is $463/tonne 

($0.21/lb). 

Table 17: Comparison of Formaldehyde Production Costs 

Process BASF Incomplete conversion & 

MeOH recovery 

Formox 

Cost of production ($/tonne) 345 364 339 

Return of capital investment ($/tonne) 33 43 48 

Cost of production with ROI ($/tonne) 378 407 387 

4.12.2 Cost of MTBE 

From 1980 to 2001, the average U.S. Gulf Coast price for MTBE fluctuated between 

$0.63/gallon (1998) and $1.27/gallon (1981) (Davenport et al. 2002). MTBE is produced 

in three different manners: (1) refinery/petrochemical plants using by-product 

isobutylene, (2) merchant plants which isomerize n- and iso-butane and dehydrogenate 

isobutane to isobutylene, and (3) tertiary butyl alcohol (TBA) plants which react byproduct TBA from propylene oxide production with methanol (Lidderdale 2001). The 

first is the cheapest way to produce MTBE at a cost of $6,000 - $10,000/BPD of capacity. 

The second is the most expensive at a cost of $20,000 - $28,000/BPD of capacity. Only 2 

plants in the U.S. use the third manner to produce MTBE. 

4.12.3 Cost of Acetic Acid 

Chemical Marketing Reporter gives the current cost (January 28th 2003) of acetic acid to 

be $992/tonne ($0.45/lb). 

4.12.4 Cost of MTG 

The capital cost of the New Zealand MTG plant (14,450 BPD) was estimated to be $767 

million dollars (mid-1980) with a total project cost of $1,475 million dollars (1985) (Bem 

1985; Maiden 1983). The production cost, including return on investment, was estimated 

to be $0.96/gallon. The New Zealand government planned to sell the gasoline for 

$1.31/gallon (Bem 1985). Initial studies performed prior to design and construction of 

the plant came up with a gasoline cost of $0.74/gallon (Bem 1985). Table 18 is a 

comparison of three other studies. 
51

Table 18: Comparison of Coal Based MTG Economic Studies 

Study (Lee et al. 1980) (Wham and Forrester 

III 1980) 

(Edwards et al. 

1979) 

Cost year 1977 1979 1978 

Coal feed rate 

(tonne/day) 

24,762 

(as received) 

18,431 

(MAF) 

18,118 

(as received) 

Coal cost $7.7/tonne $0.95/GJ $5.5/tonne 

Gasoline produced 

(BPD) 

23,065 21,921 15,000 

Price of gasoline $0.67-$0.87/gallon 

$5.1-$6.6/GJ (HHV) 

$5.5-$7.2/GJ (LHV) 

$1.25/gallon 

$9.5/GJ (HHV) 

$10.3/GJ (LHV) 

$0.90/gallon 

$6.8/GJ (HHV) 

$7.4/GJ (LHV) 

Level of detail Not very detailed – but 

refers to a more 

detailed report 

(Schreiner 1978) which 

gives overall capital 

costs and other details 

Some details – Gives total 

capital investment, 

breakdown of operating 

costs, and other economic 

assumptions 

Somewhat detailed – 

Gives capital cost by 

plant sections, 

operating costs, and 

other economic 

assumptions 

4.12.5 Cost of MTO 

One study gives the cost of methanol, olefins via MTO, and gasoline via MTG to be 

$0.45/gal ($0.68/lb), $0.16/lb, and $0.59/gal ($0.99/lb), respectively (Gradassi 1998) for 

a natural gas price of $0.47/GJ. Overall, there are not many economic details given in the 

report. 

4.12.6 Cost of DME 

At a natural gas cost of $0.6/MMBTU, DME can be produced for $0.67/gallon 

(Chaumette et al. 1999). In comparison, their estimate for methanol production at this 

natural gas price is $0.35/gallon. 

4.13 R&D Needs 

Recent research has been aimed at one-step methanol production where oxygen is reacted 

with methane to form methanol without the intermediate formation of syngas. 

Commercialization of this is believed to be at least 10 years away. 
52

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58

5 Ethanol 

5.1 Summary 

Syngas fermentation is an indirect method for producing ethanol from biomass 

feedstocks. The first step in the process is to convert biomass into a gaseous intermediate 

rich in carbon monoxide and hydrogen using gasification or other means. This gaseous 

intermediate, also known as synthesis gas, or syngas, is then converted to ethanol using 

fermentation. 

A distinct advantage of the syngas fermentation route is its ability to process nearly any 

biomass resource. Today’s corn-based ethanol industry is restricted to processing grain 

starches. Direct fermentation of biomass, as exemplified by the NREL technology, can 

handle a wider variety of biomass feedstocks, but more recalcitrant materials lead to high 

costs. Difficult-to-handle materials, softwoods for example, may best be handled with 

the syngas fermentation approach. 

Expected yields from a grassroots biomass syngas-to-ethanol facility with no external 

fuel source provided to the gasifier, are 70-105 gallons of ethanol per ton of dry biomass 

fed. The economics of this route appear to be competitive with today’s corn-based 

ethanol and projections for direct fermentation of biomass. One report states projected 

cash costs on the order of $0.70 per gallon, with feedstock available at $25 per ton. 

Capital costs are projected at about $3.00 per gallon of annual capacity. The rational 

price, defined as the ethanol sales price required for a zero net present value of a project 

with 100% equity financing and 10% real after-tax discounting, is projected to be $1.33 

per gallon. These economics would support a successful commercial project at the 

current ethanol sales price of $1.00-$1.50 per gallon. 

The syngas fermentation approach has received very modest levels of support in the past. 

Currently, there are only a handful of academic groups working in this area. More 

people, time, and dollars are needed if the pace of progress is to increase. 
59

Section Contents 

5.1 Summary ........................................................................................................................ 58

5.2 Introduction ................................................................................................................... 60

5.3 Technology Description................................................................................................. 61

5.4 Chemistry And Microbiology ........................................................................................ 63

5.5 Fermentor Design.......................................................................................................... 65

5.6 Gas Cleanliness Requirements...................................................................................... 65

5.7 Status/Technology Developers ...................................................................................... 65

5.8 Environmental Performance ......................................................................................... 66

5.9 Cost................................................................................................................................ 67

5.10 R&D Needs.................................................................................................................. 68

5.11 References.................................................................................................................... 69

List of Figures 

Figure 6: Syngas to Ethanol Simplified Block Flow Diagram .............................................. 61 

List of Tables 

Table 19: Ethanol Yields on a Common Basis....................................................................... 63 

Table 20: Summary of USDA Survey Result (Cash Costs in $/gal of ethanol)....................... 68 
60

5.2 Introduction 

Today, ethanol is the only renewable liquid fuel made in commercial quantities, with 

U.S. production in 2002 running at about 2.13 billion gal/yr, up from 1.7 billion gal/yr in 

2001 (Renewable Fuels Association 2003). According to the Energy Information 

Administration, the total U.S. demand for motor gasoline is about 8.36 million barrels per 

day, or about 1.7 GJ per year. Adjusting the figures for heating values, ethanol is 

responsible for providing only about 1% of our nation’s energy needs for gasoline type 

transportation fuels. Significant expansion of the industry is needed if ethanol is to make 

a meaningful contribution to our nation’s energy supply. The likelihood of this 

expansion is, in part, contingent on improvements in the technology used for ethanol 

production. 

Commercial production is currently by direct fermentation of carbohydrates or by 

hydration of ethylene. Direct fermentation accounts for over 95% of global capacity 

(Davenport et. al. 2002). The U.S. fermentation capacity is 2.9 billion gal/yr; non-idled 

U.S. synthetic production capacity is only 50 million gal/yr. Cornstarch hydrolyzate, 

derived from either wet milling or dry milling of corn kernels, is the major source of 

carbohydrate in the U.S. The other major ethanol producing country, Brazil - with over 

4.0 billion gal/yr of capacity, uses sugarcane juices/molasses as the major source of 

carbohydrate. 

Demand for ethanol is driven primarily by its use as a blending ingredient for gasoline. 

Non-fuel consumption in the U.S. was only 269 million gal/yr in 2001 (Davenport et. al. 

2002). Ethanol has a high octane rating, but (in many cases) meeting legislated 

requirements on gasoline oxygenate content is the market driver behind its use. The price 

of fuel ethanol traditionally tracks prices for gasoline and MTBE. MTBE is an 

alternative high-octane, gasoline-blending component made from non-renewable 

resources that can also be used to satisfy gasoline oxygenate requirements (see section 

4.10.2 for information on MTBE). 

Feedstock availability is one constraint on the ability of the existing corn-based industry 

to make a meaningful impact on our Nation’s energy supply. U.S. corn production 

reached 9.5 billion bushels in 2001 (National Corn Growers Association 2002). Fuel 

ethanol production was responsible for consuming 680 million bushels, or about 7% of 

the crop for that year. The largest consuming application for corn is direct use as animal 

feed, accounting for 5.85 billion bushels of consumption in 2001. Even if suitable 

replacements could be found for animal feed and all other uses of corn, only about 11% 

of the U.S. energy needs for gasoline type transportation fuels could be met at current 

corn production levels. 

Biomass is an abundant and low cost source of carbohydrate that often does not compete 

with animal feed or other important uses. Putsche and Sander (in Wyman 1996) estimate 

the U.S. could produce as much as 1x1012 l/yr (264 billion gal/yr) of ethanol from 

cellulosic biomasses derived from energy crops planted on idle cropland, crop residues 

such as corn stover and rice straw, municipal solid wastes, etc. 
61

One approach to utilizing this biomass resource is to modify the direct fermentation 

process so that, rather than fermenting the sugars present in starch, the sugars present in 

the cellulose and hemicellulose fractions of biomass are converted to ethanol by direct 

fermentation. Unfortunately, the cellulose and hemicellulose sugars are difficult to 

liberate and also differ in composition from the sugars present in starch. Both issues lead 

to significant differences between the two direct fermentation approaches. 

5.3 Technology Description 

For purposes of discussion in this report, the syngas intermediate required for ethanol 

production is assumed to be generated from gasification of biomass resources such as 

wood chips derived from forestry operations and other similar “low-cost” materials. It 

should be clear that other means of syngas generation could also be considered. 

Examples of these are: steam reforming of natural gas and other light hydrocarbons, 

reforming of anaerobic digester biogas, and gasification of other carbonaceous feedstocks 

such as coal, petroleum resid, coke, municipal solid waste, biomass derived fast pyrolysis 

oils, etc. 

Figure 6 is a simplified block flow diagram for a biomass syngas fermentation facility. 

The feed is first received and placed in temporary storage on-site. It is then sent to the 

gasifier where it is converted into a raw syngas mixture rich in CO and hydrogen. 

Biomass gasification has been an area of R&D interest by governments and private 

industry for many years, so several technological options exist. The syngas intermediate 

is then converted to ethanol via fermentation. Again, this approach could be applied to a 

wide variety of feedstocks. 

Figure 6: Syngas to Ethanol Simplified Block Flow Diagram 

This report only briefly discusses the indirect BCL/FERCO process since it was used in 

the design of the economic study discussed in the cost section. The BCL/FERCO gasifer 

is an indirectly heated gasifier that operates at near ambient pressure. The design 

produces a medium BTU syngas, without the need for an oxygen plant. The biomass 

feed is dried and then fed to a fast fluidized bed where it is converted into a raw syngas. 

Biomass

Syngas 

Generation, 

Conditioning, & 

Compression

Feedstock 

Handling

Syngas 

Fermentation

Ethanol 

Recovery Ethanol

Water 

and 

Solubles

Water 

and 

Nutrients

Cell 

Mass

Tail 

Gas

Ash

Biomass

Syngas 

Generation, 

Conditioning, & 

Compression

Feedstock 

Handling

Syngas 

Fermentation

Ethanol 

Recovery Ethanol

Water 

and 

Solubles

Water 

and 

Nutrients

Cell 

Mass

Tail 

Gas

Ash
62

Circulating sand provides heat for the gasification in this bed. The sand is disengaged 

along with char using a cyclone. The resulting syngas is sent downstream for further 

conditioning and compression. The disengaged sand and char is fed to a second fluidized 

bed where it is mixed with air and the char is combusted. The heat released in the 

combustor is used to heat the circulating sand. Heat is also recovered in the combustor 

offgas and used either for drying or for steam generation. 

This particular gasifier needs to operate at temperatures that are low enough that no 

slagging of the ash occurs. Typical operating temperatures and pressures are 700-850 °C 

at pressures only slightly greater than atmospheric. At these conditions, the resulting 

syngas contains significant amounts of methane, ethylene and other light hydrocarbons, 

plus tars. These species are not converted in the downstream fermentation step. 

Depending upon the design assumptions, these species can be removed in the gas 

conditioning steps, converted to additional syngas in the gas-conditioning step, or 

allowed to pass through the fermentation for eventual incineration with the fermentation 

tail gas as a fuel trim to the combustion section of the gasifier. 

The conditioned syngas is then fed to fermentation where it is converted to ethanol using 

the bacteria described in the next section. The resulting fermentation broth is quite dilute, 

typically containing 2% or less of ethanol. The ethanol can be recovered from the broth 

using recovery schemes patterned after those used in the existing corn ethanol industry 

(i.e. an ethanol-water mixture close to the azeotropic composition is distilled overhead 

and an adsorption unit is used to further dry the ethanol product to meet fuel grade 

specification on water content). The cell mass produced from the fermentation is not 

currently approved for animal feed use. A reasonable assumption for its disposal is to 

recycle it as a portion of the feed to the gasifier. 

One advantage of the syngas fermentation route is that the chemical energy stored in all 

parts of the biomass, including the lignin fraction, contributes to the yield of ethanol. The 

following equation: 

5 1.5x10

F HHVF η Gas Cond XCO H2/EtOH P + + =

 Eq. 6.1 

where 

P = Production of ethanol, million gal/yr 

F = Feed rate, tons/day (dry basis) 

HHVF = Higher heating value of the feed in Btu/lb (dry) 

ηGas+Cond = Cold gas efficiency of gasifier+conditioning steps (a fraction less than 1) 

XCO+H2/EtOH = Average conversion of CO and H2 to ethanol, as a fraction of theoretical 
63

is a simple way to generate a ballpark estimate of annual ethanol production for syngas 

fermentation. For example, assume 2,204 tons/day of dry biomass are converted with a 

higher heating value of 7,317 Btu/lb (dry), the cold gas efficiency of gasification and 

conditioning is 70%, and the average conversion of CO and H2 to ethanol is 80% of 

theoretical, then: 

P 60.2 MM gal/yr of Ethanol Production x = 5 = 1.5 10

(2204)(7317)(.7)(.8)

This equation is based on the assumption that the gasifier/condition steps converts all of 

the methane and higher hydrocarbons down to syngas and the assumption that the heating 

value of the ethanol produced is equal to the heating value of the biomass feed less 

conversion losses. It is important to include any external auxiliary fuel input to the 

gasifier (e.g. a natural gas trim to the combustor to improve control) in the F*HHVF term 

in equation 6.1 because it is based on a rough energy balance. 

Table 19 compares ethanol yields for the corn-based direct fermentation as reported in 

the USDA survey (Shapouri et. al. 2002), the biomass-based direct ferment process 

projected by NREL, and projected yields for a syngas fermentation process. It is 

somewhat surprising that yields from the syngas fermentation process are not projected to 

be higher since lignin and other components not utilized by direct methods are converted 

by the syngas fermentation route. However, this benefit appears to be largely offset by 

conversion losses in gasification, conditioning and fermentation. 

Table 19: Ethanol Yields on a Common Basis 

5.4 Chemistry And Microbiology 

The micro-organisms used for ethanol production from syngas mixtures are obligate 

anaerobes that use a heterofermentative version of the acetyl-CoA pathway for 

acetogenesis. Acetyl-CoA is produced from CO or H2/CO2 mixtures in this pathway. 

The acetyl-CoA intermediate is then converted into either acetic acid or ethanol as a 

primary metabolic product. The details of the biochemistry of acetogenesis are reviewed 

in Drake (Drake 1994). The relevant overall reactions are (ignoring cell mass 

production): 

Wet 

Mill 

Dry Mill 

Large 

Dry Mill 

Medium 

Dry Mill 

Small 

Direct 

Fermentation of 

Biomass 

Syngas 

Fermentation 

Yield, 

gal/bushel 

2.682 2.688 2.606 2.588 - - 

Yield, 

gal/ton 

(dry) 

112.7 112.9 109.5 108.7 89.7 70-105 
64

Ethanol Production: 

 ∆G’, 

 kcal/mol

HHVProducts/ 

HHVFeeds

2 3 2 2 6 CO + 3 H O → CH CH OH + 4CO -48.7 0.81 

2 CO2 + 6 H2 → CH3CH2OH + 3 H2O 28.7 0.80 

Acetic Acid Production: 

3 2 4 CO + 2 H2O → CH COOH + 2CO -39.2 0.77

2 CO2 + 4 H2 → CH3COOH + 2 H2O -25.8 0.77

The free energy changes are computed from the values listed in Table 3.7 of Roels (Roels 

1983) assuming CO2 present in the form of bicarbonate ion, liquid water and pH=7. 

Carbon monoxide is actually a preferred substrate over H2+CO2, since the change in free 

energy is more favorable. Typical CO conversions reported in the literature for 

laboratory scale fermentations are about 90%, while H2 conversions are about 70%. 

Other issues such as mass transfer between the gas and liquid phases may also contribute 

to the observed differences. 

The ratio of the heats of combustion for the products and feeds are also taken from Table 

3.7 of Roels. These values represent “cold gas efficiencies” for the fermentation step and 

are already incorporated into the denominator of Equation 6.1. It is worth noting that 

these values are rather low for an anaerobic fermentation. By comparison, the same ratio 

for production of ethanol by direct fermentation of glucose is 0.98. 

The ratio of ethanol to acetate produced is dependent upon the strain and the fermentation 

conditions. The organisms are inhibited by low pH and acetate ion concentration. When 

acetic acid is formed, the pH drops and the acetate ion concentration rises, so the 

organism switches to ethanol production to alleviate further inhibition. Typically, pH is 

kept around 4.5 in ethanol production mode. 

Many of the organisms are either mesophiles or thermophiles with temperature optimums 

ranging from room temperature to 90°C. A fairly rich media is typically, required, but 

high operating temperatures, low carbohydrate levels, low pH and high CO levels (which 

are inhibitory to methanogens) reduce the risk of contamination. 
65

5.5 Fermentor Design 

A simple gas-sparged tank reactor, operating in batch or continuous mode, can be used 

for the fermentation. While simple, this design suffers from low volumetric productivity, 

low gas conversion, and produces very dilute ethanol streams. 

Dr. Gaddy at University of Arkansas/Bioengineering Resources Inc. has studied the issue 

of fermentor design in detail. He suggests a two-stage fermentation system with cell 

recycle as a better alternative (Klasson et. al. 1991). Conditions in the first stage are 

selected to encourage cell growth, while conditions in the second stage are selected to 

encourage ethanol production. Cells are recycled in the second stage to improve 

volumetric productivity and increase conversion. 

Mass transfer between the gas and liquid phases can limit performance of syngas 

fermentation designs. Dr. Worden at Michigan State University has studied this issue in 

great detailed (Bredwell et. al. 1999). 

5.6 Gas Cleanliness Requirements 

In contrast to many other syngas based processes, syngas fermentation performance is not 

tied to a specific ratio of H2 to CO. While the organisms generally prefer CO to H2, both 

CO and H2/CO2 mixtures can be simultaneously converted. 

Very little work has been published on the effects of syngas impurities. One would 

expect some tolerance to sulfur compounds, tars and other impurities, but not enough 

work has been published to make general statements. 

5.7 Status/Technology Developers 

University of Arkansas/Bioengineering Resources Inc. - Dr. Gaddy and his associates 

began work on syngas fermentation in the late 1980s. This group has published far more 

than any other. They discovered and patented the strain Clostridium ljungdahlii (Gaddy 

and Clausen 1992), characterized its performance and developed it with the two stage 

fermentor design to the point where it was ready for scale-up from their laboratory/pilot 

plant work. Their work was supported by a U.S. DOE grant in the early 1990s under the 

concept of converting CO and H2 present from coal gasification or industrial offgases 

(i.e. refinery fuel gas, coke oven gases, offgas from carbon black manufacture, etc.) 

rather than from biomass-derived gases. Support from U.S. DOE was discontinued 

because of budget constraints in the mid 1990s. The current state of activities for this 

group is unknown, however some U.S. patents have recently been issued that include 

Celanese International Corporation as an assignee (Gaddy et. al. 2002). 

Mississippi Ethanol LLC - This company has a partially completed sawdust gasifier 

installed in Winona, MS. The facility was originally designed for methanol production, 

but market changes made these plans obsolete. The company obtained a contract with 

U.S. DOE to evaluate conversion of the existing facility to a syngas fermentation facility. 
66

The report (Mississippi Ethanol LLC 2002) discussed what would be needed for this 

conversion. The present status is unknown. 

Mississippi State University - This group, along with several other Mississippi based 

universities, is currently working under a U.S. DOE grant. Few publications have 

appeared since the work has only recently started. Dr. French presented a poster at the 

May 2003 Symposium on Biotechnology for Fuels and Chemicals in Breckenridge, CO 

titled “Isolation and Performance Optimization of Cultures Capable of Converting 

Syngas to Ethanol”. Dr. Zappi will be chairing a session on syngas fermentation at the 

November 2003 AIChE Annual Meeting. 

Oklahoma State University - This group has been working on syngas fermentation since 

the mid 1990s. Their present work is aimed at lab/pilot scale integration of gasification 

and fermentation. 

Michigan Biotechnology Institute (MBI) - Researchers from this institution published 

several articles on this subject in the early 1990s, but have since been silent. The present 

status is of this work is unknown. 

Michigan State University - Dr. Mark Worden and his associates published several 

articles on mass transfer issues for syngas fermentation in the late 1990s and early 2000s. 

Much of the work has focused on the influence of bubble size on gas-liquid mass transfer. 

Iowa State University - This group presented a poster at the May 2003 Symposium on 

Biotechnology for Fuels and Chemicals. Their focus is on production of 

polyhydroxyalkanoic acids by syngas fermentation rather than ethanol production. 

5.8 Environmental Performance 

The major driver for use of ethanol in the gasoline blending pool is as a source of 

oxygenate, in order to meet current legislative mandates. These mandates were originally 

imposed in the 1990s as a means to reduce air pollution caused by automobiles, with a 

particular target of reducing CO emissions. Opponents of the oxygenate mandate argue 

that today’s engines burn fuels more cleanly than engines of 10-15 years ago and that the 

presence of oxygenates in gasoline is no longer needed to meet CO emission targets. 

There is also some controversy over the issue of fugitive emissions from ethanol-blended 

gasolines. When ethanol is combined with other hydrocarbon blend stocks, the resulting 

blend has a vapor pressure higher than expected from only a linear combination because 

ethanol tends to form minimum boiling azeotropes with aromatics and paraffins. In the 

past, ethanol blends have received waivers on meeting vapor pressure specifications. 

MTBE is beginning to show-up as a contaminate in ground water. Leaking underground 

gasoline storage tanks is often the source. MTBE is quite mobile and persistent in the 

aquifers. Ethanol blends can leak from the same storage tanks, but ethanol is not as 

persistent, or as toxic. The foul odors, foul taste, and toxicity of MTBE have lead many 

states to propose bans on its use in gasoline. While one could argue that eliminating 
67

leaks from underground storage tanks might be a better policy direction, the need for 

oxygenate mandates has been questioned because of this issue. 

It is likely that the oxygenate mandate will be replaced with a renewable fuels standard. 

The environmental aims for this type of legislation would be to decrease net CO2

emissions by increasing the renewable fuel content of gasoline. Ethanol, the major 

renewable fuel component in today’s blend pool, will probably be used in larger volumes 

if this type of legislation were enacted. From a policy perspective, it is important to 

remember that CO2 emissions from fossil fuel consumption during the production of 

ethanol differ widely depending upon how the steam and electricity needed for the 

production process is generated. At present, the majority of corn ethanol plants consume 

fossil fuels to generate these utilities; a design that uses biomass to generate these utilities 

will result in lower overall CO2 emissions on a life cycle basis. It is unclear whether the 

upcoming renewable fuels standard will produce legislative advantages for producing 

these utilities from renewables. 

5.9 Cost 

Prices for fuel ethanol have varied between $1.00-$1.50 per gallon over the 1999-2001 

period (Davenport et. al. 2002). This price does not include excise tax exemptions. 

Table 20 presents a comparison of typical U.S. cash costs for production of ethanol from 

corn as recently reported by a USDA survey of existing plants (Shapouri et. al. 2002). 

Current corn dry mills have capital requirements of $1.50-2.00 per annual gallon of 

capacity. Cash costs vary with the type of technology used and the scale of the facility. 

Small dry mill facilities have the highest costs. These cash costs do not include interest 

expense, depreciation, income taxes or return on capital. The reported gap between 

ethanol sales price and the cash costs covers these items. Producer credits are also not 

included. The U.S. Federal government subsidizes small producers with a $0.10 per 

gallon income tax credit on up to 15 million gal/yr of production. Various states also 

have producer credit programs. 

The market price of ethanol does not reflect the actual value of ethanol in the gasoline 

blending pool. This is because of the impact of the Federal excise tax exemption. The 

Federal excise tax exemption is a benefit provided to the gasoline blender rather than the 

ethanol producer. In 2003, the excise tax exemption reduced the cost incurred by the 

gasoline blender by $0.52 per gallon of ethanol blended. This excise tax exemption will 

be reduced to $0.51 per gallon in 2005. The supporting legislation expires in 2007. The 

effect of this legislation is to lower the effective price of ethanol so that it can compete in 

the gasoline blending pool based on its usability value. The value of ethanol as a 

gasoline-blending ingredient can be estimated from its market price less the excise tax 

exemption and some corrections for logistics and blender profit margin. Under many 

market conditions, the value of ethanol is below the current cash cost of production. 

Research directed at introducing new process technologies with lower production costs is 

one means that can be pursed to remedy this situation. 
68

Table 20: Summary of USDA Survey Result (Cash Costs in $/gal of ethanol) 

Wet Mill Dry Mill 

Large 

Dry Mill 

Medium 

Dry Mill 

Small 

Net Corn 0.4795 0.5284 0.5285 0.5952 

Fuel+Power 0.1117 0.1078 0.1403 0.1786 

Other Variable 0.1861 0.1025 0.1328 0.1304 

Labor 0.0763 0.0712 0.0707 0.0962 

Other Fixed 0.0856 0.0801 0.1021 0.0877 

Total Cash Costs 0.9392 0.8900 0.9744 1.0881 

Source: Shapouri, et al. 2002 

Few studies have been published regarding ethanol production from syngas. Putsche 

(Putsche 1999) constructed an ASPEN model and estimated capital plus operating costs 

for a grassroots syngas fermentation facility that used a BCL/FERCO type of gasifier. 

Her work would best be classified as a conceptual study. Results are summarized in the 

following paragraph. 

A facility processing 2,000 tonne (dry) per day of wood would produce 48.5 million 

gal/yr of ethanol based on an ethanol yield of 71 gallons per ton. Fixed capital was 

estimated at $153.6 million, or $3.17 per annual gallon of capacity. Cash costs are 

$0.697 per gallon with feedstock cost at $25 per ton. Rational price, defined as the price 

required for a zero net present value for the project with 100% financing and 10% real 

after-tax discounting, was $1.33 per gallon. 

The Putsche results are based a reasonable set of assumptions using the current state of 

the technology. More aggressive assumptions based on future expected improvements in 

yield, reductions in capital costs, etc. would lead to more favorable economics. In any 

case, the syngas fermentation technology appears to be competitive with today’s cornbased facilities. Low cash costs are projected for the syngas fermentation facility, but 

this is offset by high capital cost. 

One of the clear advantages of the syngas fermentation approach is that just about any 

biomass resource imaginable could be processed. Present-day corn-based facilities are 

restricted to processing grain starches. Direct fermentation of biomass can handle a 

wider variety of biomass feedstocks, but more recalcitrant materials lead to much higher 

costs than the easy to process corn stover feedstock. Difficult to handle materials, 

softwoods for example, may be best handled with the syngas fermentation approach. 

5.10 R&D Needs 

The syngas fermentation approach for converting biomass resources into ethanol has 

received very modest levels of support in the past. More people, time and dollars are 

needed if the pace of progress is to increase. 
69

5.11 References 

Bredwell, M.D., Srivastava, P., Worden, R.M., “Reactor Design Issues for Synthesis-Gas 

Fermentations”, Biotechnology Progress, Vol. 15, p. 834-844, 1999. 

Davenport, R.E., Gubler, R., Yoneyama, M., Chemical Economics Handbook Marketing 

Research Report - Ethyl Alcohol, SRI International, May 2002. 

Drake, H.L. (editor), Acetogenesis, Chapman and Hall, 1994. 

Energy Information Administration, “2000 Annual Energy Review”, www.eia.doe.gov

Gaddy, J.L., Clausen, E.C., “Clostridium ljungdahlii, An Anaerobic Ethanol And Acetate 

Producing Microorganism”, US Patent 5173429, December 22, 1992. 

Gaddy, J.L., Clausen, E.C., Ko, C.W., Wade, L.E., Wikstrom, C.V., “Microbial Process 

For The Preparation Of Acetic Acid As Well As Solvent For Its Extraction From The 

Fermentation Broth”, US Patent 6368819 B1, April 9, 2002 

Klasson, K.T., Ackerson, M.D., Clausen, E.C., Gaddy, J.L., “Biorector Design for 

Synthesis Gas Fermentations”, Fuel, Vol. 70, p. 605-614, May 1991. 

Mississippi Ethanol LLC, “Final Report from Mississippi Ethanol LLC to the National 

Renewable Energy Laboratory”, NREL/SR-510-31720, March 2002. 

National Corn Growers Association, “The World of Corn 2002”, www.ncga.com

Paisely, M.A., Overend, R.P., “The SilvaGas® Process From Future Energy Resources - 

A Commercialization Success”, 12th European Conference and Technology Exhibition 

on Biomass for Energy, Industry and Climate Protection, Amsterdam, Netherlands, June 

2002, future-energy.np.def6.com/Resource.asp

Putsche, V., “Complete Process and Economic Model of Syngas Fermentation to 

Ethanol”, C Milestone Completion Report, National Renewable Energy Laboratory, 

August 31, 1999 

Renewable Fuels Association, www.ethanolrfa.org, see Ethanol Plants tab updated April 

2003. 

Roels, J.A., Energetics and Kinetics in Biotechnology, Elsevier Biomedical, 1983. 

Shapouri, H., Gallagher, P., Graboski, M.S., USDA’s 1998 Ethanol Cost-of-Production 

Survey, U.S. Department of Agriculture, Office of the Chief Economist, Office of Energy 

Policy and New Uses, Agricultural Economic Report Number 808, January 2002, 

www.usda.gov/agency/oce/oepnu/aer-808.pdf. 

Wyman, C.E. (Editor), Handbook on Bioethanol: Production and Utilization, Taylor and 

Francis, 1996. 
70

6 Mixed Higher Alcohols 

6.1 Summary 

Mixed alcohols are a more attractive gasoline blending stock for octane enhancement 

compared to methanol. Compared to methanol mixed alcohols have a lower vapor 

pressure, better solubility with hydrocarbon components, improved water tolerance, and 

higher overall heating value. There are several processes that can be used to make mixed 

alcohols from CO and H2; including isosynthesis, variants of FTS, oxosynthesis 

involving the hydroformylation of olefins, and homologation of methanol and lower 

molecular weight alcohols to make higher alcohols. 

Depending on the process conditions and catalysts used, the most abundant products are 

typically methanol and CO2. The first step in higher alcohol synthesis (HAS) is the 

formation of a C-C bond by CO insertion into CH3OH. Linear alcohols are produced in a 

stepwise fashion involving the synthesis of methanol followed by its successive 

homologation to ethanol, propanol, butanol, etc. The mechanism for HAS involves a 

complex set of numerous reactions with multiple pathways leading to a variety of 

products that are impacted by kinetic and thermodynamic constraints. The general HAS 

reaction mechanism has the following overall stoichiometry: nCO + 2nH2 → CnH2n+1OH 

+ (n-1)H2O with n typically ranging from 1 to 8 (Forzatti et al. 1991). The reaction 

stoichiometry suggests that the optimum CO/H2 = 2, however, the simultaneous 

occurrence of the WGS reaction means that the optimum ratio is closer to 1. 

Similar to other syngas conversion processes, one of the most important aspects of HAS 

is removing the large excess heat of reaction to maintain control of process temperatures, 

maximize yields, and minimize catalyst deactivation by sintering. HAS is performed in 

reactors that are similar to methanol and FT synthesis processes. In fact, to date modified 

methanol and modified FT catalysts have been the most effective in the production of 

mixed alcohols. Other catalyst types that have been researched for higher alcohol 

synthesis include sulfide-based, oxide-based, and rhodium (Rh) based. 

While other syngas-to-liquids processes have been commercialized, the commercial 

success of HAS has been limited by poor selectivity and low product yields. In general, 

advances need to be made which result in increased productivity and improved selectivity 

in higher alcohol formation. Some possible ideas include injection of lower alcohols in 

the syngas, using catalyst beds in series, and using a slurry phase reactor. 

Several studies have been performed to evaluate the economics of higher alcohol 

synthesis from natural gas. Depending of the cost of the natural gas and the economic 

assumptions used, the results of these studies varies from roughly $0.50/gallon to 

$1.20/gallon. 
71

Section Contents 

6.1 Summary ........................................................................................................................ 70

6.2 Introduction ................................................................................................................... 72

6.3 Technology Description................................................................................................. 73

6.4 Chemistry....................................................................................................................... 73

6.5 Catalysts........................................................................................................................ 76

6.6 Reactors......................................................................................................................... 79

6.7 Gas Cleanliness Requirements...................................................................................... 79

6.8 Status/Technology Developers ...................................................................................... 79

6.9 Environmental Performance ......................................................................................... 83

6.10 Cost.............................................................................................................................. 83

6.11 R&D Needs.................................................................................................................. 83

6.12 References.................................................................................................................... 85

List of Tables 

Table 21: Companies and Information Regarding Mixed Alcohols Research ...................... 82 

Table 22: EcaleneTM Composition ......................................................................................... 83 

Table 23: Comparison of Mixed Higher Alcohol Economics From Natural Gas................. 84
72

6.2 Introduction 

The production of higher alcohols from syngas has been known since the beginning of 

the last century. There are several processes that can be used to make mixed alcohols 

from CO and H2; including isosynthesis, variants of FTS, oxosynthesis involving the 

hydroformylation of olefins, and homologation of methanol and lower molecular weight 

alcohols to make higher alcohols. With the development of various gas-to-liquids 

processes (such as Fischer-Tropsch and methanol synthesis) it was recognized that higher 

alcohols were by-products of these processes when catalysts or conditions were not 

optimized. Modified FT or methanol synthesis catalysts can be promoted with alkali 

metals to shift the products towards higher alcohols. Higher alcohol synthesis (HAS) is 

also optimized at higher temperatures and lower space velocities compared to methanol 

synthesis and with H2/CO ≈ 1 instead of 2 or greater. 

While other syngas-to-liquids processes were being commercialized, the commercial 

success of HAS has been limited by poor selectivity and low product yields. Single pass 

yields of HAS are on the order of 10% syngas conversion to alcohols with methanol 

typically being the most abundant alcohol produced (Herman 2000; Wender 1996). 

Methanol can be recycled to produce more higher alcohols or removed and sold 

separately. Despite these shortcomings, in 1913 BASF patented a process to synthesize a 

mixture of alcohols, aldehydes, ketones, and other organic compounds from CO and H2

over an alkalized Co oxide catalyst at 10-20 MPa and 300-400°C (Herman 1991). 

Fischer and Tropsch developed the “Synthol” process for alcohol production in 1923. 

They used an alkalized Fe catalyst to convert syngas to alcohols at >10 MPa and 400-

450°C. Between 1935 and 1945 commercial mixed alcohol synthesis was performed 

with alkalized ZnO/Cr2O3 catalysts. The demand for mixed alcohol production from 

syngas decreased after 1945 with the increasing availability of petroleum and the desire 

for neat alcohols for manufacturing chemicals (Forzatti et al. 1991). Much of this early 

work on HAS is detailed in the review by Natta, Colombo, and Pasquon (Natta et al. 

1957). 

The oil embargo of the 1970s provided incentive for renewed interest in the synthesis and 

utilization of higher alcohols as a transportation fuel to be blended with gasoline. Mixed 

alcohols are a more attractive gasoline blending stock for octane enhancement compared 

to methanol. The octane rating is a measure of the fuel’s resistance to uncontrolled 

ignition by spurious sources such as hot spots in the combustion chamber (Mills and 

Ecklund 1989). The higher the octane number the less likelihood that a spurious ignition 

will occur. The undesirable properties of using neat methanol as a gasoline additive 

include high volatility, phase separation tendency when water is present, and 

incompatibility with certain engine fuel system components. Using mixed alcohols, 

containing methanol and higher alcohols, avoids these problems. Mixed alcohols have 

lower vapor pressure, better solubility with hydrocarbon components, improved water 

tolerance, and higher overall heating value compared to methanol. Also, when used as a 

diesel substitute at levels of 20-30 wt% higher alcohols, the calorific value, lubrication 

properties, and ignition properties are improved compared to pure methanol (Höhlein et 
73

al. 1991). Another important aspect is that it may be possible to use the existing fuel 

infrastructure. 

The environmental impact of oxygenates in gasoline is another driver which increased the 

interest in HAS. The 1990 Clean Air Act mandated the seasonal use of oxygenated 

compounds in gasoline in specific regions of the U.S. to improve air quality by reducing 

volatile organic hydrocarbon emissions that lead to increased ground level ozone 

concentrations. Until recently, the oxygenate of choice as a gasoline additive was methyl 

tert-butyl ether (MTBE). The rapid increase in demand for MTBE led to concerns about 

shortages of petroleum-derived isobutene (Verkerk et al. 1999). Commercial synthesis of 

MTBE involves the etherification of isobutene with methanol over an acidic ion 

exchange catalyst. 

Recognizing an opportunity, research efforts increased to optimize the HAS process and 

catalysts to yield mixtures of methanol and isobutanol that could be combined in a 

dehydrative coupling reaction to produce MTBE from non-petroleum sources (Burcham 

et al. 1998; Verkerk et al. 1999). The optimum molar ratio of methanol to isobutanol for 

MTBE production is one. The isobutanol is dehydrated to yield isobutene that is then 

combined with methanol in the commercial MTBE synthesis process. These activities 

may not be important in the future, considering the current environmental concerns being 

encountered with MTBE leaking into watersheds. 

6.3 Technology Description 

Currently there are no commercial plants that produce mixed alcohols in the C2 to C6

range. From a commercial viewpoint, the selectivity of HAS catalysts has not advanced 

to the point of commercial feasibility. However, the main process steps include synthesis 

gas production followed by gas clean up and conditioning, alcohol synthesis, and product 

purification. 

6.4 Chemistry 

The mechanism for HAS involves a complex set of numerous reactions with multiple 

pathways leading to a variety of products that are impacted by kinetic and 

thermodynamic constraints. No kinetic analysis of HAS has been published that is 

capable of globally predicting product compositions over ranges of operating conditions 

(Beretta et al. 1996). Depending on the process conditions and catalysts used, the most 

abundant products are typically methanol and CO2. The first step in HAS is the 

formation of a C-C bond by CO insertion into CH3OH. Linear alcohols are produced in a 

stepwise fashion involving the synthesis of methanol followed by its successive 

homologation to ethanol, propanol, butanol, etc. (Quarderer 1986). Therefore, the HAS 

catalyst should have methanol synthesis activity because methanol can be considered a 

recurrent C1 reactant. Branched higher alcohols are typically formed from modified 

methanol synthesis and modified FTS catalysts and straight chain alcohols are formed 

when alkalized MoS2 catalysts are used. 
74

The mechanism for HAS over modified high temperature methanol synthesis catalysts 

has been described as a unique carbon-chain growth mechanism that is referred to as 

oxygen retention reversal (ORR) aldol condensation with β-carbon (adjacent to the 

alcohol oxygen) addition (Herman 2000). Individual reactions in HAS can be grouped 

into several distinct reaction types (Nunan et al. 1989): 

• Linear chain growth by C1 addition at the end of the chain to yield primary linear 

alcohols. 

• Beta addition between the C1 and Cn (n ≥ 2) to yield, for example, 1-propanol and 

branched primary alcohols such as 2-methyl-1-propanol (isobutanol) for n = 2 

• Beta addition between Cm (m = 2 or 3) and Cn (n ≥ 2). 

• Methyl ester formation via carboxylic acids formed from synthesized alcohols 

• Carbonylation of methanol to yield methyl formate 

Linear alcohols can proceed along the reaction path but branched alcohols are terminal 

products of the aldol condensation pathways because they lack the 2 α-hydrogens 

required for chain growth (Hilmen et al. 1998). 

The general HAS reaction mechanism has the following overall stoichiometry (Hutchings 

et al. 1988; Wong et al. 1986): 

nCO + 2nH2 → CnH2n+1OH + (n-1)H2O (∆Hr = -61.2 kcal/mol) 

with n typically ranging from 1 to 8 (Forzatti et al. 1991). The reaction stoichiometry 

suggests that the optimum CO/H2 = 2, however, the simultaneous occurrence of the WGS 

reaction means that the optimum ratio is closer to 1. The major reactions in HAS are: 

methanol synthesis, FT reactions, higher alcohol synthesis reactions, and the water-gas 

shift (WGS) reaction (Xiaoding et al. 1987). The following is a list of some of these 

more important reactions described above that are associated with HAS. 
75

CO + 2H2 ↔ CH3OH Methanol synthesis 

CO + H2O ↔ CO2 + H2 Water-Gas Shift 

CH3OH + CO ↔ CH3CHO + H2O CO beta addition - aldehydes 

CH3OH + CO + 2H2 ↔ CH3CH2OH + H2O Ethanol Homologation 

CnH2n-1OH + CO + 2H2 ↔ CH3(CH2)nOH + H2O HAS Homologation 

2 CH3OH ↔ CH3CH2OH + H2O Condensation/dehydration 

2 CH3OH ↔ (CH3)2CO + H2O DME formation 

(CH3)2CO + H2 ↔ (CH3)2CHOH Branched iso-alcohols 

2 CH3CHO ↔ CH3COOCH2CH3 Methyl ester synthesis 

Competing reactions: 

nCO + 2nH2 → CnH2n + nH2O Olefins 

nCO + (2n+1)H2 → CnH2n+2 + nH2O Paraffins 

Methanol formation is favored at low temperatures and high pressures (Courty et al. 

1990). At high pressures, HAS increases as the temperature is increased at the expense 

of methanol formation and minimizing hydrocarbon formation. To maximize higher 

alcohols, the H2/CO ratio should be close to the usage ratio, which is about 1. Lower 

H2/CO ratios favor CO insertion and C-C chain growth. In general, the reaction 

conditions for HAS are more severe than those for methanol production. To increase the 

yield of higher alcohols, methanol can be recycled for subsequent homologation provided 

the catalyst shows good hydrocarbonylation activity (Courty et al. 1990; Quarderer 

1986). Unavoidably, the main reactions stated above produce H2O and CO2 as byproducts. WGS plays a major role and, depending on the catalyst’s shift activity, some 

chemical dehydration of alcohols can be undertaken in-situ to produce higher alcohols, 

esters, and ethers (Courty et al. 1990). Secondary reactions also produce hydrocarbons 

(HCs) including aldehydes and ketones (Courty et al. 1984; Courty et al. 1990). Also, 

frequently, substantial quantities of methane are formed (Roberts et al. 1992). 

Thermodynamic constraints limit the theoretical yield of HAS, and as in other syngas-toliquids processes, one of the most important limitations to HAS is removing the 

considerable heat of reaction to maintain control of process temperatures (Courty et al. 

1984). Compared to methanol, less alcohol product is made per mole of CO, more 

byproduct is made per mole of alcohol product, and the heat release is greater. 
76

6.5 Catalysts 

HAS catalysts are essentially bifunctional base-hydrogenation catalysts and are typically 

categorized into several groups based on their composition. Common to all HAS 

catalysts is the addition of alkali metals to the formulation. The activating character of 

alkali metal promoters is a function of their basicity. Alkali metals provide a basic site to 

catalyze the aldol condensation reaction by activating surface adsorbed CO and 

enhancing the formation of the formate intermediate. For the following, 4 groups of 

catalysts will be discussed (Herman 1991): 

1. Modified high pressure methanol synthesis catalysts – alkali-doped ZnO/Cr2O3

2. Modified low pressure methanol synthesis catalysts – alkali-doped Cu/ZnO and 

Cu/ZnO/Al2O3

3. Modified Fischer-Tropsch catalysts – alkali-doped CuO/CoO/Al2O3

4. Alkali-doped sulfides, mainly MoS2

One of the major hurdles to overcome before HAS becomes an economic commercial 

process is improved catalysts that increase the productivity and selectivity to higher 

alcohols (Fierro 1993). 

To date modified methanol and modified FT catalysts have been more effective in the 

production of mixed alcohols; the sulfide-based catalysts tend to be less active than the 

oxide-based catalysts (Herman 2000). Rhodium (Rh) based catalysts are another group 

of catalysts that are not specifically used for HAS but have been developed for selective 

ethanol synthesis. Other C2 oxygenates (i.e., acetaldehyde and acetic acid) as well as 

increased levels of methane production are also synthesized over Rh-based catalysts 

(Nirula 1994). The high cost and limited availability of rhodium for ethanol synthesis 

catalysts will impact any commercialization of these synthetic processes for converting 

syngas to ethanol (Xiaoding et al. 1987). 

6.5.1 Modified high pressure/high temperature methanol synthesis catalysts 

Alkali/ZnO/Cr2O3 - process conditions: 300-425°C and 12.5-30 MPa - major product: 

branched primary alcohols 

For non-alkalized catalysts containing Cu-Zn-Cr oxides, HAS yields were optimized for 

low Cr levels (15-21 wt% Cr). Chromia does not provide an active catalytic site but it is 

beneficial in small amounts because it acts as a structural promoter that increases the 

surface area of the catalyst and helps inhibit Cu sintering (Campos-Martin et al. 1995). 

Methanol synthesis is fast compared to HAS over a 3% K2O/ZnCr catalyst, but is still 

equilibrium limited, even at high space velocities. CO2-rich feeds have been shown to 

inhibit HAS; a 3-fold decrease in C2+ alcohol production was measured for HAS at 400°C 

with 6% CO2 (Tronconi et al. 1989). The productivity of primary alcohols was 

maximized over these alkali metal promoted high-temperature methanol synthesis 

catalysts with a CO2-free feed with H2/CO = 1 and CO conversion of ~ 5-20%. 
77

Nunan, et al. (Nunan et al. 1989) studied the effect of Cs-doping on Cu/ZnO catalysts. 

They determined that HAS product yields were maximized at 0.3-0.5 mol% Cs which 

translated to a surface concentration of 15-25%. The role of Cs-addition is to increase the 

ethanol synthesis rate that enhances the formation of higher alcohols. HAS conditions 

were 310°C at 7.6 MPa with H2/CO = 0.45. Numerous papers by Hoflund and his group 

systematically describe the properties and effectiveness of promoted Zn/Cr HAS catalysts 

(Epling et al. 1997; Epling et al. 1998; Epling et al. 1999; Hoflund et al. 1997; Hoflund et 

al. 1999; Minahan et al. 1998a; Minahan et al. 1998b). 

6.5.2 Modified low pressure/low temperature methanol synthesis catalysts 

Alkali/Cu/ZnO(Al2O3) - process conditions: 275-310°C and 5-10 MPa - major product: 

primary alcohols (Upper temperature of processes using copper catalysts is limited 

because of sintering) 

Many of the early commercial processes for HAS were based on alkali-promoted, low 

temperature methanol synthesis catalysts. Lurgi developed the Octamix process in 

collaboration with Süd Chemie, who provided the catalyst for the process. The Octamix 

catalyst contained 25-40 wt% CuO, 10-18 wt% Al2O3, 30-45 wt% ZnO, and 1.7-2.5 wt% 

K2O with a Cu:Zn ratio of 0.4-1.9 with 3-18 wt% of a variety of oxidic promoters (Cr, 

Ce, La, Mn, or Th) (Xiaoding et al. 1987). With a starting gas composition of 25-30% 

CO, 0-8% N2, 0-5% CO2, 0-5% CH4, in a balance of H2; CO conversions were between 

21-29% with 29-45% selectivity for C2+ alcohols and 17-25% CO2 selectivity at process 

conditions of 250-400°C, 10 MPa, and a gas hourly space velocity of 1000-10,000/h 

(Xiaoding et al. 1987). Methanol is the most abundant oxygenated product (~ 80%). The 

average carbon number of the oxygenated products is lower compared to products from 

modified high temperature methanol catalysts (Forzatti et al. 1991). The effectiveness of 

modified low temperature methanol synthesis catalysts has been detailed in the literature 

(Elliott and Pennella 1988; Nunan et al. 1989; Smith and Anderson 1984; Smith and 

Klier 1992). 

6.5.3 Modified Fischer-Tropsch catalysts 

Alkali/CuO/CoO-based catalysts - process conditions: 260-340°C and 6-20 MPa - major 

product: Linear primary alcohols that follow ASF distribution 

The commercial HAS process sought and developed by the Institut Français du Pétrole 

(IFP) and collaborators was based on modified Fischer-Tropsch catalysts. The IFP 

catalyst was a homogeneous mixed-oxide formulation containing Cu and Co on an 

alumina support as the active components for HAS modified with Zn and alkali metals. 

The catalysts were designed for HAS process conditions similar to low temperature 

methanol synthesis process conditions (5-15 MPa, T = 220-350°C, H2/CO = 0.5-4 with 

CO2 also as a reactant) (Xiaoding et al. 1987). Patented IFP catalyst formulations have 

the following composition, on an element basis: 10-50% Cu; 5-25% Co; 5-30% Al; 10-

70% Zn; alkali/Al = 0-0.2; Zn/Al = 0.4-2.0; Co/Al = 0.2-0.75; Cu/Al = 0.1-3.0 (Xiaoding 

et al. 1987). The homogeneity of the catalyst correlates with good catalyst performance. 
78

The IFP process yields mainly saturated, straight-chained terminal alcohols that follow an 

Anderson-Shultz-Flory-type (ASF) distribution for chain growth. At optimal conditions, 

carbon conversion efficiency of CO and CO2 is between 5 and 30% and produces a liquid 

product containing 30-50% higher alcohols with hydrocarbons being the primary 

byproducts. The lack of long-term stability and low activity of these catalysts hinders the 

commercial application of this HAS process. Catalyst lifetimes have been quoted as long 

as 8000 h at the pilot-scale with little deactivation, caused mainly by coke formation and 

sintering that decreases the homogeneity of the catalyst (Xiaoding et al. 1987). 

Additional discussions of IFP-type catalyst studies have been presented in the literature 

(Courty et al. 1984; Courty et al. 1990; Courty et al. 1982; Dai et al. 1989; Dalmon et al. 

1992). 

6.5.4 Alkali-doped sulfides, mainly MoS2

Alkali/MoS2 - process conditions: 260-350°C and 3-17.5 MPa - major product: linear 

alcohols 

In 1984, the Dow Chemical Company and Union Carbide Corporation independently 

disclosed a new catalyst system for converting syngas to linear alcohols. The catalysts 

were either supported or unsupported alkali-promoted MoS2 or Co/MoS2 (Herman 1991). 

Sulfide catalysts, such as MoS2, are well-known hydrogenation catalysts, however, 

adding alkali metals as dopants shifts the products to alcohols rather than hydrocarbons. 

The role of the alkali metal is two-fold. 1) to suppress the hydrogenation activity of the 

Mo active site and 2) to provide additional active sites for alcohol formation. Cesium is 

the most effective alkali promoter (Herman 2000), although K-addition has been 

extensively studied. The catalysts have 75-90% selectivity to higher alcohols from 

syngas with H2/CO = 1 with ~ 10% CO conversion efficiency (Herman 1991). Both 

higher alcohols and hydrocarbons formed over sufide catalysts follow a similar ASF 

molecular weight distribution. 

These catalysts have the unique property of being extremely resistant to sulfur poisoning. 

In fact, the process requires 50-100 ppm sulfur in the feed gas to maintain the sulfidity of 

the catalyst (Courty et al. 1990). //modified to achieve a more active voice// H2S in the 

feed gas is also thought to moderate the hydrogenation properties of the catalyst and 

improve selectivity of higher alcohols by reducing methanol production (Forzatti et al. 

1991). 

Sulfide catalysts are also less sensitive to CO2 in the syngas but the presence of large 

amounts (>30%) of CO2 have been shown to retard the catalyst activity (Herman 1991). 

Moderate to low (7%) CO2 concentration in the syngas does not significantly impact CO 

conversion. The selectivity to higher alcohols versus methanol is reduced by the 

presence of even low levels of CO2 (Herman 1991) so CO2 removal is recommended for 

processes using sulfide catalysts. 

Adding Co to alkalized MoS2 catalysts increases the production of ethanol and other 

higher alcohols because Co promotes the homologation of methanol to ethanol (Forzatti 

et al. 1991). Activity of the sulfide catalysts can also depend on the catalyst support 
79

material (Avila et al. 1995; Bian et al. 1998; Iranmahboob and Hill 2002; Iranmahboob et 

al. 2002; Li et al. 1998). 

6.6 Reactors 

Similar to other syngas conversion processes, one of the most important aspects of HAS 

is removing the large excess heat of reaction to maintain control of process temperatures, 

maximize yields, and minimize catalyst deactivation by sintering. HAS is performed in 

reactors that are similar to methanol and FT synthesis processes. Research and 

development is being conducted to investigate the use of slurry phase reactors for HAS. 

ChemSystems has conducted a pilot-scale study of isobutatnol synthesis in a slurry 

reactor using a Cs-promoted Cu/Zn/Al2O3 catalyst in hydrocarbon oil (40 wt% slurry) at 

12.5 MPa and 350°C (ChemSystems 1990; Herman 2000). Other HAS processes based 

on a “double bed” configuration have been explored. The idea is to optimize methanol 

production from syngas in the first reactor using a Cu-based catalyst at a lower 

temperature. The second reactor usually operates at a slightly higher temperature with a 

non-Cu Zn-chromite based catalyst to increase the yield of higher alcohols, particularly 

isobutanol by maximizing the C-C forming steps (Verkerk et al. 1999). 

6.7 Gas Cleanliness Requirements 

Gas cleanliness requirements for the modified FT and modified methanol synthesis 

catalysts are the same as those for the unmodified catalysts (see the gas cleanliness 

requirements section in the FT chapter and in the methanol chapter). Because the catalyst 

that Dow uses is already sulfided, it is extremely resistant to poisoning by sulfur in the 

syngas (1984). 

6.8 Status/Technology Developers 

As stated previously, currently there are no commercial plants that produce mixed 

alcohols in the C2 to C6 range. Table 21 gives information regarding companies that have 

actively pursued mixed alcohols research. Generally, the H2/CO ratio is 1-1.4 except for 

the Snamprogetti, which has a range of 0.5-3. In terms of commercialization, Dow, IFP 

and Snamprogetti are the most advanced in their process development (Nirula 1994). 

6.8.1 Snamprogetti (also referred to as SEHT - Snamprogetti, Enichem, and 

Haldor Topsoe) 

Snamprogetti and Haldor Topsoe developed this process jointly. A 12,000 tonne/yr pilot 

plant was started-up in 1982 in Pisticci, Italy (Olayan 1987). It was operated from 1982-

1987 (Courty et al. 1990; Mills and Ecklund 1989; Nirula 1994). The process is similar 

to that for methanol production with a different catalyst and a higher operating 

temperature. In the proposed process, syngas is produced via partial oxidation of natural 

gas. Mixed alcohols are then synthesized in fixed bed adiabatic reactors. The crude 

alcohol mixture contains 20% water. The product is purified using 3 distillation columns. 
80

The first removes methanol and ethanol, the second is a stripping column to remove 

water, and the third recovers the C3+ alcohols using azeotropic distillation with 

cyclohexane (El Sawy 1990; Fox 1993). The final water content is below 0.1 wt% (Ricci 

et al. 1984). 

The alcohol mixture was called MAS (Metanolo piu Alcoli Superiori – methanol plus 

higher alcohols) (Mills and Ecklund 1989). It was blended into gasoline at about 5 vol% 

and marketed successfully as a premium gasoline known as SUPER E (El Sawy 1990; 

Mills and Ecklund 1989). Blends of 10 vol% were also tested in 13 different types of 

European cars and compared with conventional gasoline for driveability, acceleration, 

octane, exhaust emissions, and fuel economy (Ricci et al. 1984). For all of these 

categories except emissions, the alcohol/gasoline fuel blend was shown to be similar to 

conventional gasoline. The emissions varied depending on the type of car but in general 

there was a reduction in CO (as much as 40%) and unburned HC (roughly 18%), with 

slightly higher NOx emissions (Ricci et al. 1984). Extensive research was largely 

discontinued because of the availability of large amounts of cheap petroleum. 

6.8.2 Dow 

Dow’s process was first announced in 1984. They tested 2 high pressure reactors. One 

was a fluid bed and the other was a fixed bed (Nirula 1994). Dow’s catalyst has shown 

high productivity and a resistance to sulfur. The patented Dow Chemical catalysts contain 

free or combined Mo, W, or Re, a free or combined alkali or alkaline-earth metal, and, 

optionally, a support. A catalyst containing 21% Mo and 1.5% K on a carbon support 

demonstrated 23.4% CO conversion when a H2/CO mixture equal to 0.84 was passed 

over the catalyst at 262°C and 7.2 MPa. The product distribution of the alcohols was: 25 

mol% MeOH, 12.4 mol% EtOH, 6.8 mol% 1-propanol, 2 mol% 1-butanol, and 0.3 mol% 

C5 alcohols (Quarderer and Cochran 1984). However, in addition to the mixed alcohols, 

a substantial amount of CO2, methanol, and HCs were produced (Xiaoding et al. 1987). 

The amount of methanol formed can be varied by adjusting the process conditions. The 

higher alcohols are predominantly straight chain C2-C5 alcohols. The water content is 

low (2-3 wt% given in (Quarderer 1986), and stated to be as low as 0.2% in1984 - Dow). 

6.8.3 Lurgi (also referred to as Octamix) 

While developing a low pressure methanol synthesis catalyst, a catalyst was discovered 

that had tendencies to yield higher alcohols. A 2 TPD demonstration plant was built in 

May 1990 at the Institute of Energy Process Engineering at the Research Centre Julich in 

cooperation with Lurgi (Goehna and Koenig 1989; Höhlein et al. 1991). This process 

differs from methanol synthesis only in the addition of CO2 removal and product 

purification. In Lurgi’s process scheme, syngas is produced via a combination of steam 

reforming and autothermal reforming then a tubular reactor is used for mixed alcohol 

synthesis. The water content of the crude mixture is only 1-2%. Thus a stabilizer column 

could be used instead of distillation or else molecular sieves to dry the product unless 

recovery of the methanol was desired (El Sawy 1990; Fox 1993). The final water content 
81

is low, 0.1-0.3 wt% (Goehna and Koenig 1989). The Octamix product was granted an 

EPA waver to be used as a gasoline additive on February 1, 1988 (El Sawy 1990). 

6.8.4 IFP (also referred to as Substifuel) 

IFP’s process is in the laboratory stage. They have built a 20 BPD unit in Chiba, Japan 

(Courty et al. 1990; El Sawy 1990). The IFP process scheme uses steam reforming 

followed by multibed quench synthesis reactors and distillation. The distillation section 

consists of 3 distillation columns: methanol distillation, extractive distillation with 

diethylene glycol (DEG), and distillation to recover the DEG (El Sawy 1990; Fox 1993). 

The final product contains about 0.2 wt% water. 
82

Table 21: Companies and Information Regarding Mixed Alcohols Research 

Company Country Catalyst 

type 

T (C) P (psi) H2/CO 

ratio 

Products Conversion a

Information 

Lurgi Germany Modified 

MeOH 

250-420 725-1,450 1-1.2 53.5 wt% MeOH 

41.9 wt% C2-C6 

CO conversion 

Union Carbide U.S. Rhodium 300-350 1,000-2,500 not 

found 

not found CO selectivity to 

Sagami Research Center 

Japan Rhodium 200-300 735 1.4 Mainly MeOH, 

EtOH, & CH4 

CO conversion Selectivity to etoSelectivity to alc

IFP France Modified 

MeOH 

260-320 850-1,450 1-2 30-50 wt% C2-C4 CO conversion 

Selectivity to alc

Hoechst Germany Rhodium 275 1,455 not 

found 

not found CO selectivity to 

Snamprogetti Italy Modified 

MeOH 

260-420 2,610-3,822 0.5-3 20-40 wt%C2-C4 CO conversion 

Selectivity to alc

Texaco (liquid phase system) 

U.S. Modified FT 220-240 6,615 not 

found 

12-39 wt% nonalcohol oxygenates Syngas convers

Selectivity to pro

Dow U.S. Modified FT 299-310 1,500-2,000 1.1-1.2 30-70 wt% MeOH CO conversion 

Selectivity to alc

 Source: (Courty et al. 1990; Goehna and Koenig 1989; Herman 1991; Nirula 1994; Wender 1996). 
83

6.8.5 EcaleneTM

Another mixed alcohol process, termed EcaleneTM, is currently at the bench scale but is 

being scaled up to a 500 gallon/day pilot plant. This is a collaborative project between 

Western Research Institute (WRI) and Power Energy Fuels, Inc. (PEFI), who developed 

the technology. Table 22 gives the composition of their alcohol mixture which can vary 

depending on the process parameters. The main alcohol constituent is ethanol. The 

synthesis reactor operates at 290-360°C and 145-1,595 psi (Lucero et al. 2001). 

Table 22: EcaleneTM Composition 

Constituent Current range (wt%) 

(Taylor 2002) 

Target composition (wt%) 

(Lucero et al. 2001) 

Methanol 5-30 0 

Ethanol 45-75 75 

Propanol 15 9 

Butanol 5 7 

Pentanol 3 5 

Hexanol & higher 2 4 

6.9 Environmental Performance 

The evaporative emissions from fuel blends containing mixed alcohols (particularly 

methanol and ethanol) are typically much higher than emissions from straight gasoline. 

As a fuel additive, mixed alcohols have been shown to exhibit a reduction in CO, HCs, 

and nitric oxide emissions (Höhlein et al. 1991). 

6.10 Cost 

Table 23 summarizes the economics of higher alcohol synthesis from several studies 

which examine natural gas as the feedstock. Another source states a breakdown of the 

capital cost to be 50% for syngas production, 29% for alcohol synthesis, 17% for CO2

removal, and 4% for alcohols fractionation (Courty et al. 1990). Economies of scale 

would improve the economics of mixed alcohol synthesis but it is also extremely 

important to reduce energy losses as well as overall heat and momentum transfer duty 

(Lange and Tijm 1996). 

6.11 R&D Needs 

In general, advances need to be made which result in increased productivity and 

improved selectivity in higher alcohol formation. Some possible ideas include injection 

of lower alcohols in the syngas, using catalyst beds in series, and using a slurry phase 

reactor (Herman, 2000). 
84

Table 23: Comparison of Mixed Higher Alcohol Economics From Natural Gas 

Study (ChemSystems 

1990) 

(Salmon 1986) (Courty et al. 

1984) 

(Ricci et al. 1984) 

(El Sawy 1990) 

Cost year 1986 1989 1983 1982 1986 Natural gas cost ($/GJ) 

$1.9/GJ and $0.47/GJ 

$1.9/GJ and $0.47/GJ 

$4.3/GJ $0.8-$6.6/GJ $2.8/GJ, $2.6/GJ an

$0.47/GJ 

Methanol/higher alcohol mixture (wt%) 

50/50 65/35 

for 70-90% selectivity 

(60-70)/(30-40) 70/30 70/30 and 50/50 

Mixed alcohols produced (tonne/yr) 

550,000 386,000 600,000 730,000 635,000 – 730,000 

Price of mixed alcohols 

$167.7/tonne $0.505/gallon $105.6/tonne $0.318/gallon 

Production cost: 

($1.9/GJ nat gas) $198-$255/tonne $0.60-$0.77/gallon ($0.47/GJ nat gas) $162-$199/tonne $0.49-$0.60/gallon 

Production cost: 

$313.9/tonne $0.95/gallon With 15% return: 

$400/tonne $1.21/gallon Production cost: 

For $1/GJ nat gas $150/tonne $0.45/gallon 

$110-$395/tonne $0.33-$1.19/gallon 

Production cost: 

($2.8/GJ nat gas) $282/tonne $0.774/gallon ($0.47/GJ nat gas) $205/tonne $0.562/gallon ($2.6/GJ nat gas) $0.90-$1.04/gallon 

Level of detail Few details Few details Few details Few details Few details 

Note: Production cost = no return on investment for these references. 
85

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Campos-Martin, J. M., Guerreroruiz, A., and Fierro, J. L. G. (1995). "Structural and 

Surface Properties of CuO-ZnO-Cr2O3 Catalysts and Their Relationship with Selectivity 

to Higher Alcohol Synthesis." Journal of Catalysis, 156(2), 208 - 218. 

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Configurations of Integrated-Gasification-Combined-Cycle/Once-Through." EPRI/GS6869. 

Courty, P., Arlie, J. P., Convers, A., Mikitenko, P., and Sugier, A. (1984). "C1-C6 

Alcohols from Syngas." Hydrocarbon Processing, 63(11), 105-108. 

Courty, P., Chaumette, P., Raimbault, C., and Travers, P. (1990). "Production of 

methanol-higher alcohol mixtures from natural gas via syngas chemistry." Revue de 

l'Institut Francais du Petrole, 45(4), 561-78. 

Courty, P., Durand, D., Freund, E., and Sugier, A. (1982). "C1-C6 alcohols from 

synthesis gas on copper-cobalt catalysts." J. Mol. Catal., 17(2-3), 241-54. 

Dai, L., Chen, Z., Li, G., Li, Y., and Liu, X. (1989). "Synthesis of C1-C5 alcohols from 

coal-based syngas." Proceedings - Annual International Pittsburgh Coal Conference, 

6th(2), 739-46. 
86

Dalmon, J. A., Chaumette, P., and Mirodatos, C. (1992). "Higher alcohols synthesis on 

cobalt based model catalysts." Catalysis Today, 15(1), 101-127. 

Elliott, D. J., and Pennella, F. (1988). "Mechanism of ethanol formation from synthesis 

gas over copper oxide/zinc oxide/alumina." Journal of Catalysis, 114(1), 90-9. 

El Sawy, A. H. (1990). "Evaluation of Mixed Alcohol Production Processes and 

Catalysts." NTIS. DE90010325. SAND89-7151, Mitre Corporation. 

Epling, W. S., Hoflund, G. B., Hart, W. M., and Minahan, D. M. (1997). "Reaction and 

surface characterization study of higher alcohol synthesis catalysts, II. Cs-promoted 

commercial Zn/Cr spinel." Journal of Catalysis, 172(1), 13-23. 

Epling, W. S., Hoflund, G. B., and Minahan, D. M. (1998). "Reaction and surface 

characterization study of higher alcohol synthesis catalysts. VII. Cs- and Pd-promoted 1: 

1 Zn/Cr spinel." Journal of Catalysis, 175(2), 175-184. 

Epling, W. S., Hoflund, G. B., and Minahan, D. M. (1999). "Higher alcohol synthesis 

reaction study. VI: effect of Cr replacement by Mn on the performance of Cs- and Cs, Pdpromoted Zn/Cr spinel catalysts." Applied Catalysis, A: General, 183(2), 335-343. 

Fierro, J. L. G. (1993). "Catalysis in C1 chemistry: future and prospect." Catalysis 

Letters, 22(1-2), 67-91. 

Forzatti, P., Tronconi, E., and Pasquon, I. (1991). "Higher Alcohol Synthesis." Catalysis 

Reviews-Science and Engineering, 33(1-2), 109-168. 

Fox, J. M., III. (1993). "The different catalytic routes for methane valorization: an 

assessment of processes for liquid fuels." Catalysis Reviews - Science and Engineering, 

35(2), 169-212. 

Goehna, H., and Koenig, P. (1989). "The Octamix process." Indirect Liquefaction, Proc., 

Contract. Rev. Meet.(CONF-891131--DE90 008422), 59-83. 

Herman, R. G. (1991). "Chapter 7 - Classical and Non-Classical Routes for Alcohol 

Synthesis." New Trends in CO Activation, L. Guczi, ed., Elsevier, New York, 265-349. 

Herman, R. G. (2000). "Advances in catalytic synthesis and utilization of higher 

alcohols." Catalysis Today, 55(3), 233-245. 

Hilmen, A. M., Xu, M., Gines, M. J. L., and Iglesia, E. (1998). "Synthesis of higher 

alcohols on copper catalysts supported on alkali-promoted basic oxides." Applied 

Catalysis A: General, 169(2), 355-372. 

Hoflund, G. B., Epling, W. S., and Minahan, D. M. (1997). "Higher alcohol synthesis 

reaction study using K-promoted ZnO catalysts. III." Catalysis Letters, 45(1,2), 135-138. 
87

Hoflund, G. B., Epling, W. S., and Minahan, D. M. (1999). "An efficient catalyst for the 

production of isobutanol and methanol from syngas. XI. K- and Pd-promoted Zn/Cr/Mn 

spinel (excess ZnO)." Catalysis Letters, 62(2-4), 169-173. 

Höhlein, B., Mausbeck, D., Supp, E., and König, P. "Selective Catalysis for the Synthesis 

of Energy Alcohols Status and Outlook." International Symposium on Alcohol Fuels, 

Florence, Italy, 43-48. 

Hutchings, G. J., Copperthwaite, R. G., and Coville, N. J. (1988). "Catalysis for 

hydrocarbon formation and transformations." South African Journal of Science, 84(1), 

12-16. 

Iranmahboob, J., and Hill, D. O. (2002). "Alcohol synthesis from syngas over 

K2CO3/CoS/MoS2 on activated carbon." Catalysis Letters, 78(1-4), 49-55. 

Iranmahboob, J., Toghiani, H., Hill, D. O., and Nadim, F. (2002). "The influence of clay 

on K2CO3/Co-MoS2 catalyst in the production of higher alcohol fuel." Fuel Processing 

Technology, 79(1), 71-75. 

Lange, J. P., and Tijm, P. J. A. (1996). "Processes for converting methane to liquid fuels: 

economic screening through energy management." Chemical Engineering Science, 

51(10), 2379-2387. 

Li, X., Feng, L., Liu, Z., Zhong, B., Dadyburjor, D. B., and Kugler, E. L. (1998). "Higher 

Alcohols from Synthesis Gas Using Carbon-Supported Doped Molybdenum-Based 

Catalysts." Industrial & Engineering Chemistry Research, 37(10), 3853-3863. 

Lucero, A. J., Klepper, R. E., O'Keefe, W. M., and Sethi, V. K. (2001). "Development of 

a process for production of mixed alcohols from synthesis gas." Preprints of Symposia - 

American Chemical Society, Division of Fuel Chemistry, 46(2), 413-419. 

Mills, G. A., and Ecklund, E. E. (1989). "Alternative Fuels - Progress and Prospects .2." 

Chemtech, 19(10), 626-631. 

Minahan, D. M., Epling, W. S., and Hoflund, G. B. (1998a). "An efficient catalyst for the 

production of isobutanol and methanol from syngas. VIII: Cs- and Pd-promoted Zn/Cr 

spinel (excess ZnO)." Catalysis Letters, 50(3,4), 199-203. 

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reaction study. V. Effect of excess ZnO on catalyst performance." Applied Catalysis, A: 

General, 166(2), 375-385. 

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Alcohols from Carbon Monoxide and Hydrogen." Catalysis, P. H. Emmett, ed., Reinhold, 

New York, 131-174. 

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88

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(1989). "Higher alcohol and oxygenate synthesis over cesium-doped copper/zinc oxide 

catalysts." Journal of Catalysis, 116(1), 195-221. 

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Saudi Arabia." Energy Progress, 7(1), 9-17. 

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New Orleans, Louisiana. 

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alcohols from hydrogen and carbon monoxide." PCT Int. Appl., (Dow Chemical Co., 

USA). Wo, 40 pp. 

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methanol and higher alcohols from synthesis gas." Chemia Stosowana, 28(1), 155-68. 

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synthesis." J. Catal., 85(2), 428-36. 

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Abstracts of Papers of the American Chemical Society, 203, 82-PETR. 

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synthesis gas Utilizing Catalytic Thermal Reforming (CTR) and the "Power Energy 

Process". 

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Carbon Oxides and Hydrogen .17. - Higher Alcohol Synthesis over Alkali MetalPromoted High- Temperature Methanol Catalysts." Applied Catalysis, 47(2), 317-333. 

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89

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Catalysis Today, 2(1), 125-170. 
90

7 Fischer-Tropsch Synthesis 

7.1 Summary 

Two main characteristics of Fischer-Tropsch synthesis (FTS) are the unavoidable 

production of a wide range of hydrocarbon products (olefins, paraffins, and oxygenated 

products) and the liberation of a large amount of heat from the highly exothermic 

synthesis reactions. Product distributions are influenced by temperature, feed gas 

composition (H2/CO), pressure, catalyst type, and catalyst composition. FT products are 

produced in four main steps: syngas generation, gas purification, FT synthesis, and 

product upgrading. Depending on the types and quantities of FT products desired, either 

low (200–240°C) or high temperature (300–350°C) synthesis is used with either an iron 

(Fe) or cobalt catalyst (Co). 

FTS catalysts can lose activity as a result of 1) conversion of the active metal site to an 

inactive oxide site, 2) sintering, 3) loss of active area by carbon deposition, and 4) 

chemical poisoning. Some of these mechanisms are unavoidable and others can be 

prevented or minimized by controling the impurity levels in the syngas. By far the most 

abundant, important, and most studied FTS catalyst poison is sulfur. Other catalyst 

poisons include halides and nitrogen compounds (e.g., NH3, NOx and HCN). 

Over the years, four types of FT reactors have been designed and used commercially. 

The fixed bed tubular reactor known as the ARGE reactor operates at 220-260°C and 20-

30 bar. High temperature circulating fluidized bed reactors, known as Synthol reactors, 

have been developed for gasoline and light olefin production. They operate at 350°C and 

25 bar. Recently, the Sasol Advanced Synthol reactor has been developed. It is a fixed 

fluidized bed reactor with similar operating conditions as the Synthol reactor at half the 

capital cost and size for the same capacity. The fourth reactor design is the low 

temperature slurry reactor which is a 3-phase reactor consisting of a solid catalyst 

suspended and dispersed in a high thermal capacity liquid (often the FT wax product). 

Syngas is bubbled through the liquid phase achieving excellent contact with the catalyst 

while keeping the catalyst particles dispersed. Slurry reactors are optimized at low 

temperatures for FT high wax production with low methane production. 

The first FT plants began operation in Germany in 1938 but closed down after the second 

world war. Then in 1955, Sasol, a world-leader in the commercial production of liquid 

fuels and chemicals from coal and crude oil, started Sasol I in Sasolburg, South Africa. 

Following the success of Sasol I, Sasol II and III came on line in 1980 and 1982, 

respectively. The syngas at these three plants as well as at several other plants abroad is 

converted to more than 200 fuel and chemical products. In the early 1990s, two other FT 

plants came on line. The Mossgas plant which converts natural gas to FT products using 

a high temperature process and an iron catalyst started up in South Africa in 1992. 

Additionally, Shell commissioned a plant in 1993 in Bintuli, Malaysia using the Shell 

Middle Distillate Synthesis process, which is essentially enhanced FT synthesis. 

Currently, Syntroleum is building a 10,000 BPD specialty chemicals and lube oil plant 

located in Northwestern Australia. 
91

Section Contents 

7.1 Summary ........................................................................................................................ 90

7.2 Introduction ................................................................................................................... 92

7.3 Technology Description................................................................................................. 93

7.4 Chemistry....................................................................................................................... 94

7.5 Catalysts........................................................................................................................ 96

7.6 Reactors......................................................................................................................... 97

7.7 Gas Cleanliness Requirements...................................................................................... 99

7.8 Status/Technology Developers .................................................................................... 101

7.9 Process Performance................................................................................................... 102

7.10 Environmental Performance...................................................................................... 102

7.11 Cost............................................................................................................................ 103

7.12 R&D Needs................................................................................................................ 106

7.13 References.................................................................................................................. 107

List of Figures 

Figure 7: FTS General Process Flow Diagram .................................................................... 93 

Figure 8: Anderson-Shulz-Flory Distribution ....................................................................... 95 

Figure 9: Types of Fischer-Tropsch Synthesis Reactors....................................................... 98 

List of Tables 

Table 24: Syngas Impurities and Tolerances for Fischer-Tropsch ..................................... 100 

Table 25: Price Ranges for Natural Gas, Gasoline, and Diesel ......................................... 104 

Table 26: Comparison of Biomass Based Fischer Tropsch Studies.................................... 106 
92

7.2 Introduction 

The synthesis of hydrocarbons from CO hydrogenation over transition metal catalysts 

was discovered in 1902 when Sabatier and Sanderens produced CH4 from H2 and CO 

mixtures passed over Ni, Fe, and Co catalysts. In 1923, Fischer and Tropsch reported the 

use of alkalized Fe catalysts to produce liquid hydrocarbons rich in oxygenated 

compounds — termed the Synthol process. Succeeding these initial discoveries, 

considerable effort went into developing catalysts for this process. A precipitated Co 

catalyst with 100 parts by weight Co, 5 parts by weight ThO2, 8 parts by weight MgO, 

and 200 parts by weight kieselguhr (silicious diatomatous earth) became known as the 

“standard” atmospheric pressure process catalyst. In 1936, Fischer and Pilcher developed 

the medium pressure (10-15 bar) FTS process. Following this development, alkalized Fe 

catalysts were implemented into the medium pressure FTS process. Collectively, the 

process of converting CO and H2 mixtures to liquid hydrocarbons over a transition metal 

catalyst has become known as the Fischer-Tropsch synthesis. 

Two main characteristics of Fischer-Tropsch synthesis (FTS) are the unavoidable 

production of a wide range of hydrocarbon products and the liberation of a large amount 

of heat from the highly exothermic synthesis reactions. Consequently, reactor design and 

process development has focused heavily on heat removal and temperature control. The 

focus of catalyst development is on improved catalyst lifetimes, activity, and selectivity. 

Single pass FTS always produces a wide range of olefins, paraffins, and oxygenated 

products such as alcohols, aldehydes, acids and ketones with water as a byproduct. 

Product distributions are influenced by temperature, feed gas composition (H2/CO), 

pressure, catalyst type, and catalyst composition. Product selectivity can also be 

improved using multiple step processes to upgrade the FTS products. 

In a review by Frohning, et al. (1954), it was cited that upwards of 4000 publications and 

a similar number of patents dealing with FTS could be found in the literature. Since then 

FTS has attracted an enormous amount of research and development effort. A 

comprehensive bibliography of FTS literature, including journal and conference articles, 

books, government reports and patents can be found in the Fischer-Tropsch Archive at 

www.fischer-topsch.org. This website is sponsored by Syntroleum Corporation in 

cooperation with Dr. Anthony Stranges, Professor of History at Texas A&M University 

and contains more than 7500 references and citations. This site has collected a 

bibliography of the large body of documents from the 1920's through the 1970s, which 

are important for researching the history and development of FTS and related processes 

as well as an up-to-date listing of the latest publications in this field. Many excellent 

reviews of FTS have been drawn upon for this report (Mills, 1993; Dry, 2002; and Dry, 

1981) in an attempt to summarize, the chemistry (Frohning, et al, 1982), catalyst 

development (Bartholemew, 1991; Oukaci, et al, 1999; and Raje, et al, 1997), 

commercial processes (Dry, 2002, 1999, 1982; Senden, et al, 1992; Haid, et al, 2002, 

Van Nierop et al, 2000), reactor development (Adesini, 1986; Dry, 1988; and Dry and 

Hoogendoorm, 1981), and economics of FTS. 
93

7.3 Technology Description 

There are four main steps to producing FT products: syngas generation, gas purification, 

FT synthesis, and product upgrading. Figure 7 depicts a generic process flow diagram. 

When using natural gas as the feedstock, many authors (Abbott and Crewdson, 2002; 

Rostrup-Nielsen, 2000; Vosloo, 2001; and Wilhelm, et al, 2001) have recommended 

autothermal reforming or autothermal reforming in combination with steam reforming as 

the best option for syngas generation. This is primarily attributed to the resulting H2/CO 

ratio and the fact that there is a more favorable economy of scale for air separation units 

than for tubular reactors (steam methane reforming - SMR). If the feedstock is coal, the 

syngas is produced via high temperature gasification in the presence of oxygen and 

steam. Depending on the types and quantities of FT products desired, either low (200–

240°C) or high temperature (300–350°C) synthesis is used with either an iron or cobalt 

catalyst. FTS temperatures are usually kept below 400°C to minimize CH4 production. 

Generally, cobalt catalysts are only used at low temperatures. This is because at higher 

temperatures, a significant amount of methane is produced. Low temperatures yield high 

molecular mass linear waxes while high temperatures produce gasoline and low 

molecular weight olefins. If maximizing the gasoline product fraction, it is best to use an 

iron catalyst at a high temperature in a fixed fluid bed reactor. If maximizing the diesel 

product fraction, a slurry reactor with a cobalt catalyst is the best choice. The FT reactors 

are operated at pressures ranging from 10-40 bar (145–580 psi). Upgrading usually 

means a combination of hydrotreating, hydrocracking, and hydroisomerization in 

addition to product separation. 

Figure 7: FTS General Process Flow Diagram 

Coal, 

Natural 

Gas, or 

Biomass

Gas Cleanup 

and Conditioning Gasifier

Slurry (Co) or 

Tubular (Fe) 

Reactor

Diesel

Clean syngas

H2 and CO

•Particulate Removal

•Wet Scrubbing 

•Catalytic Tar Conversion

•Sulfur Scrubbing

•WGS, etc.

Air, Oxygen, 

Steam

Waxes (>C20)

CFB or FFB (Fe) 

Reactor

Low T FTS

High T FTS

Hydrocracking

Gasoline

Olefins

(C3 – C11)

Oligomerization

Isomerization

Hydrogenation

Coal, 

Natural 

Gas, or 

Biomass

Gas Cleanup 

and Conditioning Gasifier

Slurry (Co) or 

Tubular (Fe) 

Reactor

Diesel

Clean syngas

H2 and CO

•Particulate Removal

•Wet Scrubbing 

•Catalytic Tar Conversion

•Sulfur Scrubbing

•WGS, etc.

Air, Oxygen, 

Steam

Waxes (>C20)

CFB or FFB (Fe) 

Reactor

Low T FTS

High T FTS

Hydrocracking

Gasoline

Olefins

(C3 – C11)

Oligomerization

Isomerization

Hydrogenation
94

7.4 Chemistry 

FTS has long been recognized as a polymerization reaction with the basic steps of: 

1. reactant (CO) adsorption on the catalyst surface 

2. chain initiation by CO dissociation followed by hydrogenation 

3. chain growth by insertion of additional CO molecules followed by hydrogenation 

4. chain termination 

5. product desorption from the catalyst surface 

Chemisorbed methyl species are formed by dissociation of absorbed CO molecules and 

stepwise addition of hydrogen atoms. These methyl species can further hydrogenate to 

form methane or act as initiators for chain growth. Chain growth occurs via sequential 

addition of CH2 groups while the growing alkyl chain remains chemisorbed to the metal 

surface at the terminal methylene group. Chain termination can occur at any time during 

the chain growth process to yield either an α-olefin or an n-paraffin once the product 

desorbs. 

The following is the FTS reaction: 

1. CO + 2H2 Æ --CH2-- + H2O ∆Hr (227°C) = -165kJ/mol (Haid, et al, 2000) 

The water-gas shift (WGS) reaction is a secondary reaction that readily occurs when Fe 

catalysts are used. Combining reaction one (above) with reaction two (below) gives the 

net reaction for Fe catalyzed FTS (reaction 3). 

2. CO + H2O Æ H2 + CO2 (Water-Gas Shift)

3. 2CO + H2 Æ --CH2-- + CO2 (net overall FTS) 

The required H2 to CO ratio for the cobalt catalyst is 2.15 but since the iron catalyst 

performs WGS in addition to the FT reaction, the H2 to CO ratio can be slightly lower for 

the iron catalyst, 1.7 (Dry, 2002). 

Specific FTS products are synthesized according to the following reactions. 

4. CO + 3H2 Æ CH4 + H2O (Methanation) 

5. nCO + (2n+1)H2 Æ CnH2n+2 + nH2O (Paraffins) 

6. nCO + 2nH2 Æ CnH2n + nH2O (Olefins) 

7. nCO + 2nH2 Æ CnH2n+1OH + (n-1)H2O (Alcohols) 
95

Another competing reaction that becomes important in FTS is the Boudouard reaction: 

8. 2CO Æ Cs + CO2 

Carbon deposition on the catalyst surface causes catalyst deactivation. 

FTS is kinetically controlled and the intrinsic kinetics is stepwise chain growth, in effect 

the polymerization of CH2 groups on a catalyst surface. FTS product selectivities are 

determined by the ability of a catalyst to catalyze chain propagation versus chain 

termination reactions. The polymerization rates, and therefore kinetics, are independent 

of the products formed. The probability of chain growth and chain termination is 

independent of chain length. Therefore, selectivities of various hydrocarbons can be 

predicted based on simple statistical distributions calculated from chain growth 

probability and carbon number. The chain polymerization kinetics model known as the 

Anderson-Shulz-Flory (ASF) model is represented by the following equation: 

Wn = n(1-α)

2

α

n-1

Wn is the weight percent of a product containing n carbon atoms and α is the chain 

growth probability. This equation is graphically represented in Figure 8. It clearly 

displays the predicted distributions for several products and product ranges of particular 

interest. 

Figure 8: Anderson-Shulz-Flory Distribution

Probability of Chain Growth (α)

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

Weight Fraction

0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

CH4

C2

C3

Gasoline

(C5-C11)

Diesel

(C9-C25)

Wax

(C35-C120)
96

Irrespective of operating conditions, the FT reaction always produces a range of olefins, 

paraffins, and oxygenated compounds (alcohols, aldehydes, acids, and ketones). There 

are several unique aspects of FT products. Regardless of the product type, they are 

predominantly linear with high olefinicity. In fact, the paraffin-to-olefin ratio is lower 

than thermodynamically predicted. The olefins that do form are predominantly terminal 

(alpha). A considerable amount of monomethyl chain branches form and the degree of 

branching decreases as the chain length increases. Theoretically, only methane can be 

produced with 100% selectivity (Senden, et al, 1992). The only other product that can be 

produced with high selectivity is heavy paraffin wax. The gasoline product fraction has a 

maximum selectivity of 48%. The maximum diesel product fraction selectivity is closer 

to 40% and varies depending on the range of carbon numbers in the product cut. The 

variables that influence the product distribution are: reactor temperature, pressure, feed 

gas composition, catalyst type, and promoters. 

7.5 Catalysts 

Group VIII transition metal oxides are generally regarded as good CO hydrogenation 

catalysts. The earliest catalysts used for FTS were Fe and Co. Vannice has reported a 

relative activity of these metals for FTS (Adesina, 1996). In decreasing order of activity: 

Ru > Fe > Ni > Co > Rh > Pd > Pt. Ni is basically a methanation catalyst and does not 

have the broad selectivity of other FT catalysts. Ru has very high activity and quite high 

selectivity for producing high molecular weight products at low temperatures. Fe is also 

very active and has WGS activity. Fe readily forms carbides, nitrides, and carbonitrides 

with metallic character that also have FTS activity. Fe also has a stronger tendency than 

Ni or Co to produce carbon that deposits on the surface and deactivates the catalyst. Co 

tends to have a longer lifetime than Fe catalysts and does not have WGS activity, which 

leads to improved carbon conversion to products because CO2 is not formed. Co 

catalysts in FTS yield mainly straight chain hydrocarbons (no oxygenates like Fe). 

Although Ru is the most active FTS catalyst it is 3x105 times more expensive than Fe. 

Iron is by far the least expensive FTS catalyst of all of these metals. Co catalysts are 230 

times more expensive than Fe but are still an alternative to Fe catalysts for FTS because 

they demonstrate activity at lower synthesis pressures, so higher catalyst costs can be 

offset by lower operating costs. 

The three key properties of FT catalysts are lifetime, activity, and product selectivity. 

Optimizing these properties for desired commercial application has been the focus of FT 

catalyst research and development since the processes were first discovered. Each one of 

these properties can be affected by a variety of strategies including; 

• use of promoters (chemical and structural) 

• catalyst preparation and formulation 

• pretreatment and reduction 

• selective poisoning 

• shape selectivity with zeolites 
97

The performance of Co catalysts is not very sensitive to the addition of promoters. Early 

work demonstrated that the addition of ThO2 improved wax production at atmospheric 

pressure, but had little effect at higher pressures. With Fe catalysts, however, promoters 

and supports are essential catalyst components. Since the discovery of FTS, potassium 

has been used as a promoter for Fe catalysts to effectively increase the basicity of the 

catalyst surface. The objective is to increase the adsorption of CO to the metal surface, 

which tends to withdraw electrons from the metal, by providing an electron donor. 

Adding Potassium oxide to Fe catalysts also tends to decrease hydrogenation of adsorbed 

carbon species, so chain growth is enhanced, resulting in a higher molecular weight 

product distribution that is more olefinic. Potassium promotion also tends to increase 

WGS activity leading to a faster rate of catalyst deactivation because of the increased rate 

of carbon deposition of the surface of the catalyst. 

Copper has also been successfully used as a promoter in Fe FTS catalysts. It increases 

the rate of FTS more effectively than potassium, but decreases the rate of the WGS 

reaction. Copper has been shown to facilitate iron reduction. The average molecular 

weight of the products increases in the presence of copper, but not as much as when 

potassium is used. Potassium promotion is not effective for Co catalysts. 

Catalyst preparation impacts the performance of Fe and Co catalysts. Fe catalysts can be 

prepared by precipitation onto catalyst supports such as SiO2 or Al2O3, or as fused iron 

where formulations are prepared in molten iron, then cooled and crushed. The role of 

supports in Co catalysts is also important. Since Co is more expensive than Fe, 

precipitating the ideal concentration of metal onto a support can help reduce catalyst 

costs while maximizing activity and durability. 

The combination of light transition metal oxides such as MnO with Fe increases the 

selectivity of light olefins in FTS. Fe/Mn/K catalysts have shown selectivity for C2-C4

olefins as high as 85-90%. Noble metal addition to Co catalysts increases FTS activity 

but not selectivity. 

7.6 Reactors 

One of the challenges with FTS, mentioned above, is the removal of the large amount of 

excess heat generated by the exothermic synthesis reactions. Insufficient heat removal 

leads to localized overheating which results in high carbon deposition leading to catalyst 

deactivation. Methane formation also dominates at higher temperatures at the expense of 

desired FTS products. For large-scale commercial FTS reactors heat removal and 

temperature control are the most important design features to obtain optimum product 

selectivity and long catalyst lifetimes. Over the years, basically four FTS reactor designs 

have been used commercially. Figure 9 depicts the types of Fischer-Tropsch synthesis 

reactors. 
98

Figure 9: Types of Fischer-Tropsch Synthesis Reactors 

One of the earliest FTS reactor designs was the fixed-bed tubular reactor. After many 

years of development, Ruhrchemie and Lurgi have refined this concept into what is 

known as the ARGE high capacity FT reactor. This has been used since the inception of 

Sasol I. The reactors contain 2,000 tubes filled with Fe catalyst immersed in boiling 

water for heat removal. The water bath temperature is maintained in the reactor by 

controlling the pressure. Syngas is introduced into the top of the reactor, flows through 

the tubes, and the products exit at the bottom of the reactor. Wax accounts for 50% of the 

products. Conversion efficiencies are on the order of 70%. The reactor is operated at 20-

30 bar at an operating temperature of 220-260oC. Additional temperature control is 

obtained by using high gas velocities and gas recycling. Recycled gas is typically 2.5 

times the amount of fresh syngas introduced into the reactor. Catalyst lifetimes are 

around 70-100 days (Wender, 1996) and catalysts removal can be quite difficult. 

High-temperature circulating fluidized-bed reactors have been developed for gasoline and 

light olefin production. These reactors are known as Synthol reactors and operate at 

350°C and 25 bar. The combined gas feed (fresh and recycled) enters at the bottom of 

the reactor and entrains catalyst that is flowing down the standpipe and through the slide 

valve. The high gas velocity carries the entrained catalyst into the reaction zone where 

heat is removed through heat exchangers. Product gases and catalyst are then transported 

into a large diameter catalyst hopper where the catalyst settles out and the product gases 

exit through a cyclone. These Synthol reactors have been successfully used for many 

years, however, they have a number of limitations. They are physically very complex 

reactors that involve circulation of large amounts of catalyst that leads to considerable 

erosion in particular regions of the reactor. Synthol reactors have also reached their 

design limit in terms of capacity and throughput (7,500 BPD capacity, (Lutz, 2001)). 

Catalyst

Tubes

Gas Outlet

Wax Outlet

Syngas Intlet

Multi-tubular

(ARGE)

Fixed Bed

Gas

Outlet

Hopper

Catalyst

Syngas

Inlet

Standpipe

Circulating

(Synthol)

Fluidized Bed

Fixed

(Sasol Advanced Synthol)

Fluidized Bed

Gas

Outlet

Syngas

Inlet

Catalyst

Bed

Cooling

Water

Steam

Gas

Outlet

Syngas

Inlet

Cooling

Water

Steam

Wax

Out

Slurry

Bed

Fixed

Slurry Bed
99

The fixed fluidized bed Sasol Advanced Synthol reactor has replaced the circulating 

fluidized bed Synthol reactor. Gas is introduced through a distributor and bubbles up 

through the catalyst bed. Heat is removed by an internal heat exchanger immersed in the 

catalyst bed. These new reactors are half the cost and size of the circulating reactors for 

the same capacity. They also have better thermal efficiency with a less severe 

temperature gradient and a lower pressure drop across the reactor. Operating costs are 

considerably lower and there is greater process flexibility (in terms of product 

distribution) and the possibility for scale-up to 20,000 BPD (Lutz, 2001). Process 

conditions in the fixed fluidized bed reactors are similar to those established in the 

Synthol reactors. 

Another reactor design is the low-temperature slurry reactor, which is a design that has 

been considered since Kolbel’s pioneering work in the 1950s (Dry, 1996; and Dry, 2002). 

These 3-phase reactors consist of a solid catalyst suspended and dispersed in a high 

thermal capacity liquid (often the FT wax product). Syngas is bubbled through the liquid 

phase achieving excellent contact with catalyst while keeping the catalyst particles 

dispersed. Slurry reactors are optimized at low temperatures for high FT wax production 

with low methane production. Compared to the fluidized bed reactors, the liquid slurry 

bed offers the advantages of better temperature control, lower catalyst loading, and 

significantly lower catalyst attrition rates. The improved isothermal conditions in slurry 

bed reactors allows for higher average reactor temperatures leading to higher conversions 

to products. Slurry bed reactors also cost 75% less than the much more complex 

multitubular fixed bed reactors. These reactors have only recently been put into 

commercial FT production primarily because one of the technical barriers, which 

required considerable development, was reliable catalyst separation from the FT waxes. 

7.7 Gas Cleanliness Requirements 

FTS catalysts can lose activity as a result of 1) conversion of the active metal site to an 

inactive oxide site, 2) sintering, 3) loss of active area by carbon deposition, and 4) 

chemical poisoning. Some of these mechanisms are unavoidable and others can be 

prevented or minimized by insuring that the impurity levels in the incident syngas are 

acceptable for the given process. 

Carbon deposition is the most important mode of catalyst deactivation that can be 

impacted by addition of promoters to catalysts and reaction temperature and pressure. 

This mode of catalyst deactivation is largely unavoidable and FTS processes must be 

operated in a manner that the decreasing output from coke deposition is balanced with the 

economic considerations of catalyst regeneration or replacement. In general, because of 

its high activity, the coke deposition rate is higher for Fe catalyst than Co catalyst. 

Consequently, Co catalysts have longer lifetimes. 

One of the more controllable modes of catalyst deactivation is that induced by poisoning 

of the active sites by impurities in the syngas. By far the most abundant, important, and 

most studied FTS catalyst poison is sulfur. Sulfur is present in both natural gas and coal 

and during steam reforming or gasification gets converted primarily to H2S plus other 
100

organic sulfur compounds. Sulfur compounds rapidly deactivate both iron and cobalt 

catalysts, presumably by forming surface metal sulfides that do not have FTS activity. 

Ideally, there should be no sulfur in the syngas. There is, however, always a small 

amount that gets through to the catalyst. There is really no safe sulfur level in FTS. 

Again, the level of gas cleaning required is based on economic considerations; namely 

how long the catalyst remains active versus the investment in gas cleaning. 

In the beginning, Fischer recommended 4 ppm as the maximum sulfur content in the FTS 

feedgas. Dry (1981) recommends a maximum sulfur content of 0.2 ppm based on 

commercial experience in the Sasol plants. Co catalysts are more sensitive to sulfur 

poisoning than Fe catalysts. Given the relative cost of Co versus Fe, more efficient sulfur 

removal should be expected for FTS with Co catalysts. Several more recent references 

cite even lower sulfur tolerances. Boerrigter, et al, (2002) state that sulfur levels in 

syngas for FTS should be below 1 ppmv and Turk, et al, (2001) claim that the total sulfur 

content in syngas should be 60 ppb. In general, for a fixed bed reactor design, any 

catalyst poison will have the most pronounced affect near the gas inlet and propagate 

through the reactor towards the outlet, whereas in a fluidized bed design the poison will 

have a uniform affect throughout the reactor. 

Other syngas impurities are also known to poison FTS catalysts. Halide levels in syngas 

should be less than 10 ppb (Boerrigter, et al, 2002) and referenced nitrogen levels are 10 

ppmv NH3, 0.1 ppmv NOx and 10 ppb HCN (Turk, et al, 2001). Additionally, water 

oxidizes FT catalysts (both Fe & Co) but the rate of oxidation is higher for the iron 

catalyst. Plus water has an inhibiting effect on the iron catalyst because of its WGS 

activity (Espinoza, et al, 2000). Table 24 summarizes the syngas impurities and 

tolerances. 

Table 24: Syngas Impurities and Tolerances for Fischer-Tropsch 

Impurity Tolerance Level Source 

Sulfur 0.2 ppm 

1 ppmv 

60 ppb 

Dry, 1981 

Boerrigter, et al, 2002 

Turk, et al, 2001 

Halides 10 ppb Boerrigter, et al, 2002 

Nitrogen 10 ppmv NH3 

0.2 ppmv NOx 

10 ppb HCN 

Turk, et al, 2001 

Note: There are differing points of view regarding sulfur level but in general, 

the sulfur content of the syngas should be minimized according to economics. 

Commercial processes are available to clean syngas to meet these stringent contaminant 

requirements. The Rectisol process uses chilled methanol to scrub the raw syngas. NH3, 

H2S, tars, and CO2 are removed from syngas to required levels. Other chemical 

absorption processes include potassium carbonate or alkanolamine (MEA–

monoethanolamine or DEA–diethanolamine) for wet scrubbing. Fixed bed reactors 

containing ZnO are also used for sulfur polishing. Whether or not these gas-cleaning 

processes are economical will depend on the scale of the FTS process. 
101

7.8 Status/Technology Developers 

The first FT plants began operation in Germany in 1938. There were nine plants with a 

total capacity of 660,000 tonnes/yr of FT products. These plants closed down after the 

second world war (Dry, 2002). In 1955, Sasol, a world-leader in the commercial 

production of liquid fuels and chemicals from coal and crude oil, started up it’s first FT 

plant (Sasol I) in Sasolburg, South Africa. The plant’s capacity is 6 million tonnes/yr of 

FT products from coal. There are five tubular fixed bed (ARGE) reactors for wax 

production that are still operational today and three circulating fluid bed (CFB) reactors 

(Synthol reactors), which operated until 1993 (Dry, 1982). In 1993, a slurry reactor that 

matches the production of the five ARGE reactors began operation. Following the 

success of Sasol I, Sasol II and III came on line in 1980 and 1982, respectively. These 

plants are located in Secunda, South Africa. Sasol II and III all have circulating fluid bed 

(CFB) reactors (Synthol reactors) but with improved heat exchange. Thus, they get a 

three-fold increase in capacity compared to the synthol reactors at Sasol I. From 1995 to 

1999, these second generation CFB reactors (Synthol reactors; 16 of them) were replaced 

with eight fixed fluid bed (FFB) reactors known as Sasol Advanced Synthol reactors 

(Dry, 2002). This was done for economic reasons and although the plant capacity did not 

increase, the operating cost was projected to decrease by $1/BBL. In general, these 

reactors are advantageous because of lower capital costs, higher efficiency, lower 

operating costs, and less downtime for maintenance. CFB reactors suffered, initially, 

from enormous scale-up problems (mechanical, operational, & catalyst point of view). 

Sasol’s South Africa plants use both low and high temperature FT synthesis. The coal 

consumption at Sasol’s three commercial plants is 36 million tons/year and that 

generates, on average, 1.5 million Nm3

/hr of syngas from 97 pressurized Lurgi coal 

gasifiers (Van Nierop, et al, 2000). The syngas is converted to "more than 200 fuel and 

chemical products at its plants in Sasolburg and Secunda in South Africa, as well as at 

several other plants abroad." (http://www.gasandoil.com/goc/company/cna02527.htm) 

Its products; including gasoline, diesel, candle waxes, hard waxes, hydrocarbon 

lubricants, methane, phenol and cresol, tar and pitch, ammonia, detergents, etc., are 

exported to more than 70 countries around the world. Sasol also supplies 41% of South 

Africa’s liquid transportation fuel requirements. The company has developed worldleading technology for the commercial production of synthetic fuels and chemicals from 

low-grade coal as well as the conversion of natural gas to environment-friendly diesel 

and chemicals. 

In the early 1990s, two other FT plants came on line. The Mossgas plant, which converts 

natural gas to FT products using a high temperature process and an iron catalyst, started 

up in South Africa in 1992 (Dry, 2002). This plant produces 1 million tonnes/yr of FT 

products including motor gasoline, distillates, kerosene, alcohols and LPG. Additionally, 

Shell commissioned a plant in 1993 in Bintuli, Malaysia, using the Shell Middle 

Distillate Synthesis process, which is essentially enhanced FT synthesis. This plant 

produces 500,000 tonnes/yr (12,000 BPD) of FT products from 100 MMSCFD of natural 

gas using a cobalt catalyst (Senden, et al, 1992). The plant produces automotive fuels, 

specialty chemicals, and waxes. 
102

Currently, Syntroleum is building a 10,000 BPD plant (specialty chemicals and lube oil) 

in the Burrup Peninsula in Northwestern Australia (Haid, et al, 2000). The engineering 

began in 1999. Construction started in 2001 and operation is expected to begin in 2003. 

The plant will convert natural gas in the gas-to-liquids and specialty products plant using 

their proprietary process and a cobalt catalyst. 

In early 2000, Sasol studied the feasibility of replacing coal with recently discovered 

offshore natural gas. They decided that switching to natural gas would reduce 

expenditures in its coal mining operations, as well as the high costs of compliance with 

environmental regulations associated with coal. Therefore, they are beginning to build a 

network of natural gas pipelines, which are expected to deliver gas to their facilities in 

2004. Sasol plans to use the natural gas as a supplementary feedstock at Secunda and as 

a replacement feedstock for the coal in Sasolburg (http://www.oilbarrel.com/archives/features_archive/2002/jan-2002/sasol310102.htm and 

http://www.eia.doe.gov/emeu/cabs/safrica.html). 

7.9 Process Performance 

In reforming methane, about 20% of the carbon is converted to CO2 while, for coal, this 

number is 50% due to the lower hydrogen content of coal (Dry, 2002). Coal gasification 

produces a syngas with a H2:CO ratio of about 0.67. For this ratio, the maximum 

obtainable conversion of CO to FT products without WGS is 33% (Raje, et al, 1997). 

The theoretical max conversion of methane to paraffins is 78% on a LHV basis. The 

conversion at the Bintuli plant in Malaysia is about 63% (Senden, et al, 1992). In a slurry 

reactor, very small catalyst particles are suspended in a liquid medium allowing effective 

heat removal and good temperature control. Sasol’s slurry phase process has a thermal 

efficiency of about 60% and a carbon conversion efficiency of about 75% (Lutz, 2001). 

The slurry reactor has several advantages over the multitubular design. For example, the 

capital cost is 25% less for the same capacity, the pressure difference across the reactor is 

less (making the gas compression cost lower), there is a four fold lower catalyst 

consumption rate, and the reactor can operate at higher temperatures and thus higher 

conversions (Dry, 1982). One disadvantage is that any poison will poison the entire 

amount of catalyst, whereas, in the fixed tube design, only the top portion of the catalyst 

will be poisoned. 

7.10 Environmental Performance 

Compared to conventional fuels, FT fuels contain no sulfur and low aromatics. These 

properties, along with a high cetane number, result in superior combustion characteristics 

(Jager, 1998 and Alleman and McCormick, 2003). Tests performed on heavy duty trucks 

showed decreases in vehicle emissions of HC, CO, NOx, and PM when using a FT fuel 

(Haid, et al, 2000; Lutz, 2001). FT diesel has been tested in a variety of light- and heavyduty vehicles and engines. The paper by Alleman and McCormick (2003) summarizes 

FT diesel fuel property and emission information found in the literature. Overall, FT 
103

diesel showed a reduction in regulated as well as some unregulated emissions compared 

to conventional diesel. 

Several life cycle assessments (LCA) have been performed on a variety of transportation 

fuels including FT diesel and gasoline. Most of the studies have examined only 

greenhouse gas emissions and energy consumption (General Motors, 2001; Marano and 

Cieferno, 2001; and MIT, 2000) with the exception of General Motors. They also 

examined five criteria pollutants (VOCs, CO, NOx, PM10, and SOx). These analyses 

have examined the emissions and energy consumption from resource extraction to end 

use. The results of the studies vary based on the feedstock procurement, technology 

conversion, and vehicle assumptions. However, in general, there is not a big advantage 

to FT liquids from fossil fuels in terms of energy consumption and green house gas 

emissions. This will not be the case for biomass systems. Because of the improved 

combustion characteristics of the FT liquids, performing a complete LCA including 

criteria pollutants will mostly likely show the overall benefits of FT liquids compared to 

conventional transportation fuels. 

7.11 Cost 

Table 25 presents information about the price of natural gas, gasoline, and diesel fuel, 

prior to giving cost information about FT synthesis. Over the years the price of each of 

these has fluctuated significantly. 
104

Table 25: Price Ranges for Natural Gas, Gasoline, and Diesel 

Fuel Average 

U.S. price 

(1996-2002) 

High Low Current price 

(Nov 02 avg) 

Natural yearly avg monthly avg yearly avg monthly avg 

gas 

$2.76/Mscf 

(national 

wellhead) 

$4.12/Mscf

$3.8/GJ 

(2001) 

$8.06/Mscf 

$7.4/GJ 

(Jan 01) 

$1.96/Mscf 

$1.8/GJ 

(1998) 

$2.35/Mscf 

$2.2/GJ 

(Jan 02) 

$3.64/Mscf 

$3.4/GJ 

Gasoline $1.33/gallon 

(includes 

taxes) 

$1.82/gallon 

$14/GJ (HHV) 

$15/GJ (LHV) 

(May 2001) 

$0.97/gallon 

$7/GJ (HHV) 

$8/GJ (LHV) 

(Feb. 1999) 

$1.42/gallon 

($11/GJ, HHV; 

$12/GJ, LHV) 

Excluding dist., 

marketing & taxes:

$0.74/gallon 

($5.6/GJ, HHV; 

$6.1, LHV) 

Diesel $1.25/gallon 

(includes 

taxes) 

$1.67/gallon 

$11/GJ (HHV) 

$12/GJ (LHV) 

(Oct. 2000) 

$0.95/gallon 

$6/GJ (HHV) 

$7/GJ (LHV) 

(Feb. 1999) 

$1.42/gallon 

($9.7/GJ, HHV; 

$10.5, LHV) 

Excluding dist., 

marketing & taxes:

$0.75/gallon 

($5.1/GJ, HHV, 

$5.5, LHV) 

Data from EIA: 

(1) U.S. Natural Gas Prices and Natural Gas Summary from the Short-Term Energy 

Outlook (http://www.eia.doe.gov/oil_gas/natural_gas/info_glance/prices.html ) (Note: 

For natural gas, $/Mscf is essential equivalent to $/MMBtu). Note: Mscf=1,000 standard 

cubic feet 

(2) Gasoline and Diesel Fuel Update (http://tonto.eia.doe.gov/oog/info/gdu/gasdiesel.asp ) 

The following is cost information specific to the existing commercial plants and studies 

by others for fossil fuel feedstocks. In general, the publications give facts without a lot of 

details. 

− The capital cost of the Bintuli, Malaysia plant was about $660 million. (Senden, et al, 

1992) 

− The capital cost of replacing Sasol’s 16 CFB reactors with 8 FFB reactors was $225 

million. There was no increase in capacity but the reduction in the operating cost was 

projected to be $1/BBL. (Chang, 2000) 

− Studies by Sasol indicate that a two train GTL plant made up of slurry phase reactors 

producing 30,000 BPD of liquid transport fuels can be constructed at a capital cost of 

about $25,000/BPD capacity including utilities, offsite facilities, and infrastructure. 

(Lutz, 2001 and Vosloo, 2001) 

− Sasol examined the economics of a grass root plant using slurry phase reactors 

feeding 100 MMSCFD of natural gas and producing 10,000 BBL/day (425,000 
105

ton/yr) of transportation fuel. The capital investment was around $300 million 

resulting in an investment of $30,000/daily BBL. (Jager, 1998) 

− At a natural gas price of $0.50/MMBtu (inexpensive), the feedstock accounts for 

$5/BBL of the product price and the operating cost (fixed and variable) is estimated 

to contribute an additional $5/BBL to the product price. (Gradassi, 1998; Jager, 

1998; Lutz, 2001; and Vosloo, 2001) 

− Using literature information, an economic analysis was performed for a 50,000 BPD 

plant using a feedstock cost of roughly $0.50/MMBtu (inexpensive). The resulting 

cost of the FT liquids at an internal rate of return of 15% was $26/BBL 

($0.83/gallon). This means that in order for the natural gas to FT liquids process to 

be competitive at this low feedstock cost, the price of crude oil needs to be at least 

$18/BBL. (Gradassi, 1998) 

− FT liquids from both coal and natural gas were examined for a plant that produces 

roughly 50,000 BPD of FT liquids. At a natural gas cost of $0.5/MMBtu 

(inexpensive) the cost of the FT liquids is $24/BBL ($0.76/gallon) and at $4/MMBtu 

this cost increases to $52/BBL ($1.7/gallon). Using Illionois #6 coal, the cost of the 

FT liquids is $46/BBL ($1.5/gallon). For a once-through plant that co-produces 

electricity which is sold at $0.05/kWh, the cost of the FT is reduced to $35/BBL 

($1.1/gallon). (Gray and Tomlinson, 1997) 

The majority of the capital cost for FT production, greater than 50%, comes from syngas 

generation (Dry 2002; Senden, et al, 1992, and Vosloo, 2001). Additionally, the cost of 

syngas production from methane is 30% lower than that from coal and it is a more 

efficient process (Dry 2002). Any cost reduction in syngas generation will have a large 

impact on the economics. 

Three studies that have examined the production of FT liquids from biomass are 

summarized in Table 26. 
106

Table 26: Comparison of Biomass Based Fischer Tropsch Studies 

Study Tijmensen (2000) (a), (b) Mitre (1996) (a), (b) Novem (2000) (a), (c)

Biomass feed rate 

(BD tonne/day) 

1,920 2,000 1,358 

Biomass cost $2/GJ 

$38/dry tonne 

$2.45/GJ 

$46/dry tonne 

$3/GJ 

$55/dry tonne 

Electricity selling 

price 

$0.057/kWh $0.05/kWh $0.067/kWh 

Net power (MW) 80 – 100 from biomass 

(150 MW gas turbine) 

110 – 120 about 150 

Other fuels co-fires nat gas in gas 

turbine 

None co-fires nat gas in gas 

turbine 

Raw or finished 

products 

raw Finished finished 

FT fuels produced 1,216 - 2,026 BBL/day 1,357 - 1,715 BBL/day 486 - 1,378 BBL/day 

Cost year 1999 1993 1999 

Price of products 

(without distribution 

costs and taxes) 

$13 - $30/GJ 

$1.8 - $4.1/gallon 

base case = $15 - 16/GJ 

$8 - $14/GJ 

$1.1 - $1.9/gallon 

$9 - $13/GJ 

$1.2 - $1.8/gallon 

Level of detail Very detailed - Gives costs 

of individual equipment, 

operating costs, and other 

economic parameters and 

assumptions 

Somewhat detailed - 

Gives costs of major plant 

sections, operating costs, 

and lists other economic 

assumptions 

Some details - Gives 

costs of major plant 

sections, and 

assumptions used to 

determine installed 

costs 

Notes: (a) Examined direct & indirect gasifiers and atmospheric and pressurized gasifiers. 

(b) Examined both maximum liquids (recycle of unconverted syngas) and once 

through with power generated from unconverted syngas. 

(c) Only examined once through with combined heat and power. 

7.12 R&D Needs 

In as much as syngas generation accounts for the largest capital cost, efficient least-cost 

biomass gasification is essential for the conversion of biomass to FT products. It will be 

important to develop gas clean up that is economical at various scales of biomass to FT 

conversion. Additionally, selectivity and catalyst lifetime will play a role in the success 

of FT products from biomass. 
107

7.13 References 

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process." Oil & Gas Journal 100(16): 64-66, 68, 70. 

Adesina, A.A. (1996). "Hydrocarbon synthesis via Fischer-Tropsch reaction: travails and 

triumphs", Appl. Catal. A, 13: 345-367. 

Alleman, T.L. and McCormick, R.L. (2003). "Fischer-Tropsch Diesel Fuels - Properties 

and Exhaust Emissions: A Literature Review". 2003 SAE World Congress, CI Engine 

Combustion Processes & Performance with Alternative Fuels (SP-1737), Detroit, 

Michigan, March 3-6. 

Bartholomew, C. H. (1991). "Recent Developments in Fischer-Tropsch Catalysis," 

Chapter 5 in "New Trends in CO Activation", (L. Guczi ed.), Stud. Surf. Sci. Catal., 64: 

158. 

Boerrigter, H., den Uil, H., and Calis, H.-P. (2002). "Green Diesel from Biomass via 

Fischer-Tropsch Synthesis: New Insights in Gas Cleaning and Process Design." Paper 

presented a Pyrolysis and Gasification of Biomass and waste, Expert Meeting, 30 

September, 2002, Strasbourg, FR. 

Chang, T. (2000). "South African company commercializes new F-T process." Oil & Gas 

Journal 98(2): 42-45. 

Davis, B. H. (2001). "Fischer-Tropsch synthesis: current mechanism and futuristic 

needs," Fuel Proc. Tech., 71(1-3): 157-166. 

Dry, M. E. (2002). "The Fischer-Tropsch process: 1950-2000." Catalysis Today 71(3-4): 

227-241. 

Dry, M. E. (1999). "Fischer-Tropsch reactions and the environment", Appl. Catal. A, 

189(2): 185-190. 

Dry, M.E. (1996). "Practical and Theoretical Aspects of the Catalytic Fischer-Tropsch 

Process," Appl. Catal. A, 138: 319 

Dry, M. E. (1990). "The Fischer-Tropsch process-commercial aspects", Cat. Today, 6: 

183. 

Dry, M. E. (1988). "The Sasol route to chemicals and fuels." Studies in Surface Science 

and Catalysis 36(Methane Convers.): 447-56. 

Dry, M. E. (1982). "Catalytic Aspects of Industrial Fischer-Tropsch Sunthesis." Journal 

of Molecular Catalysis 71: 133-144. 

Dry, M.E. (1981). "The Fischer-Tropsch Synthesis," in Catalysis, Science and 

Technology, (J.R. Anderson, M. Boudart, eds.), Springer-Verlag, Vol. 1, pp. 159-256. 
108

Dry, M. E. and Hoogendoorn, J.C. (1981). "Technology of the Fischer-Tropsch process." 

Catal. Rev. - Sci. Eng. 23(2): 265-78. 

Espinoza, R. L., Du Toit, E., Santamaria, J., Menendez, M., Coronas, J. and Irusta, S. 

(2000). "Use of membranes in Fischer-Tropsch reactors." Studies in Surface Science and 

Catalysis 130A(International Congress on Catalysis, 2000, Pt. A): 389-394. 

Espinoza, R. L., Steynberg, A. P., Jager, B., and Vosloo, A. C. (1999). "Low temperature 

Fischer-Tropsch synthesis from a Sasol perspective", Appl. Catal. A, 186(1-2): 13-26. 

Frohning, C., Kolbel, H., Ralek, M., Rottig, W., Schnur, F., and Schulz, H. (1982). 

"Fischer-tropsch process", Chapter 8 in Chemical feedstocks from coal, J. Falbe, ed., 

John Wiley and Sons, NY, pp. 309-432. 

General Motors, Argonne National Laboratory, ExxonMobil, Shell. (2001). "Well-toWheel energy Use and Greenhouse Gas emissions of Advanced Fuel/Vechile Systems - 

North American Analysis." 

Gradassi, M. J. (1998). "Economics of gas to liquids manufacture." Studies in Surface 

Science and Catalysis 119(Natural Gas Conversion V): 35-44. 

Gray, D. and Tomlinson, G. (1997). "A technical and economic comparison of natural 

gas and coal feedstocks for Fischer-Tropsch synthesis." Studies in Surface Science and 

Catalysis 107(Natural Gas Conversion IV): 145-150. 

Haid, M. O., Schubert, P. F. and Bayens, C. A. (2000). "Synthetic fuel and lubricants 

production using gas-to-liquids technology." DGMK Tagungsbericht 2000-3: 205-212. 

http://www.eia.doe.gov/emeu/cabs/safrica.html 

http://www.oil-barrel.com/archives/features_archive/2002/jan-2002/sasol310102.htm 

Jager, B. (1998). "Developments in Fischer-Tropsch technology." Studies in Surface 

Science and Catalysis 119(Natural Gas Conversion V): 25-34. 

Jager, B., and Espinoza, R. L. (1995). "Advances in low temperature Fischer-Tropsch 

synthesis", Cat. Today, 23: 17. 

Jess, A., Popp, R., and Hedden, K. (1999). "Fischer-Tropsch-synthesis with nitrogen-rich 

syngas: Fundamentals and reactor design aspects", Appl. Catal. A, 186(1-2): 321-342. 

Kolbel, H., Ralek, M. (1980). "Fischer-Tropsch synthesis in the liquid phase", Catal. Rev. 

- Sci. Eng., 21(2): 225-274. 

Lutz, B. (2001). "New age gas-to-liquids processing." Hydrocarbon Engineering 6(11): 

23-26,28. 
109

Marano, J. J., and Ciferno, J. P. (2001). "Life-Cycle Greenhouse-Gas Emissions 

Inventory For Fischer-Tropsch Fuels". U.S. DOE. National Energy Technology 

Laboratroy. 

Massachusetts Institute of Technology Energy Laboratory. (2000). "On the Road in 

2020: A Life-Cycle Analysis of New Automobile Technologies". MIT EL 00-003. 

Mills, G. (1993) Status and future Opportunities for Conversion of Synthesis Gas to 

Liquid Energy Fuels: Final Report. National Renewable Energy Laboratory, Golden, 

CO, TP-421-5150. 

MITRE Corporation. (1996). Techno-Economic Assessment of Biomass Gasification 

Technologies for Fuels and Power. 

Novem. (2000). "Technical and Economic Data Biomass-Bases Energy Conversion 

Systems for the Production of Gaseous and/or Liquid Energy Carriers". Report No. 

GAVE-9915. 

Oukaci, R., Singleton, A. H., and Goodwin, J. G. (1999). "Comparison of Patented Co FT Catalysts using Fixed-Bed and Slurry Bubble Column Reactors", Appl. Catal. A, 186: 

129-144. 

Patzlaff, J., Liu, Y., Graffmann, C., and Gaube, J. (1999). "Studies on product 

distributions of iron and cobalt catalyzed Fischer-Tropsch synthesis", Appl. Catal. A, 

186(1-2): 109-119. 

Raje, A., Inga, J. R. and Davis, B. H. (1997). "Fischer-Tropsch synthesis: process 

considerations based on performance of iron-based catalysts." Fuel 76(3): 273-280. 

Rao, V. U. S, Stiegel, G. J., Cinquegrane, G. J., and Srivastave, R. D. (1992). "Iron based 

catalysts for slurry phase Fischer-Tropsch process: Technology Review", Fuel Proc. 

Tech., 30: 83-107. 

Rostrup-Nielsen, J. R. (2000). "New Aspects of Syngas Production and Use." Catalysis 

Today 63: 159-167. 

Schulz, H. (1999). "Short history and present trends of Fischer-Tropsch synthesis", Appl. 

Catal. A, 186(1-2): 3-12. 

Schulz, H. (1985). "Selectivity and mechanism of the Fischer-Tropsch COhydrogenation", C1 Mol. Chem., 1: 231 

Senden, M. M. G., Sie, S. T., Post, M. F. M. and Ansorge, J. (1992). "Engineering 

aspects of the conversion of natural gas into middle distillates." NATO ASI Series, Series 

E: Applied Sciences 225(Chemical Reactor Technology for Environmentally Safe 

Reactors and Products): 227-47. 
110

Steynberg, A. P., Espinoza, R. L., Jager, B., and Vosloo, A. C. (1999). "High temperature 

Fischer-Tropsch synthesis in commercial practice", Appl. Catal. A, 186(1-2): 41-54. 

Tijmensen, M. J. A., Faaij, A. P. C., Hamelinck, C. N. and van Hardeveld, M. R. M. 

(2002). "Exploration of the possibilities for production of Fischer Tropsch liquids and 

power via biomass gasification." Biomass & Bioenergy 23(2): 129-152. 

Tijmensen, M. J. A., Faaij, A. P. C., Hamelinck, C. N. and van Hardeveld, M. R. M. 

(2000) "The production of Fischer Tropsch liquids and power through biomass 

gasification", M. Sc. Thesis, Utrecht University/Science, Technology and Society, 

Utrecht. Report No. NWS-E-2000-29. 

Turk, B.S., Merkel, T., Lopez-Ortiz, A., Gupta, R.P. Portzer, J.W., Kishnam, G., 

Freeman, B.D., and Fleming, G.K. (2001). "Novel TEchnologies for Gaseous 

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September 2001. 

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20th Century as a Building Block for the 21st." Presented at the 2000 Gasification 

Technologies Council, October 8-11, 2000. San Francisco, CA 

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Monoxide and Hydrogen", Catal. Rev.-Sci. Eng., 14(2): 153-191. 

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297. 
111

8 Oxosynthesis 

8.1 Summary 

The oxosynthesis process is also known as hydroformylation. It involves the reaction of 

CO and H2 with olefinic hydrocarbons to form an isomeric mixture of normal- and isoaldehydes. It is an industrial synthetic route for the conversion of olefins (in the C3-C15

range) to produce solvents, synthetic detergents, flavorings, perfumes and other 

healthcare products, and other high value commodity chemicals. Worldwide production 

of oxo-aldehydes and alcohols was 6.5 million tons per year in 1997. 

Oxosynthesis is a rapid reaction catalyzed by soluble cobalt or rhodium complexes. 

Three complimentary catalytic hydroformylation processes have been developed and 

commercialized. Phosphine-modified rhodium catalysts are best to produce 2-

ethylhexanol, a plasticizer alcohol used to make flexible PVC, from propylene and 

syngas. This is the highest volume oxosynthesis process. Unmodified cobalt catalysts 

are versatile enough to be used for hydroformylation of high carbon number mixed 

olefins. Phosphine-modified cobalt catalysts are used for the production of higher, 

detergent range alcohols rather than aldehydes. The LP OxoTM Low Pressure Oxo 

Process is the world's leading process for the production of oxo alcohols from olefins. 
112

Section Contents 

8.1 Summary ...................................................................................................................... 111 

8.2 Introduction ................................................................................................................. 113

8.3 Technology Description............................................................................................... 113

8.4 Chemistry..................................................................................................................... 114

8.5 Catalysts...................................................................................................................... 115

8.6 Gas Cleanliness Requirements.................................................................................... 116

8.7 Status/Technology Developers .................................................................................... 116

8.8 Cost.............................................................................................................................. 119

8.9 R&D Needs.................................................................................................................. 119

8.10 References.................................................................................................................. 120

List of Figures 

Figure 10: Schematic Flow Diagram of the LP Oxo Process............................................. 117 

List of Tables 

Table 27: Summary of U.S. Oxo Chemical Producers ........................................................ 118 

Table 28: U.S. Prices for Oxo Chemicals............................................................................ 119 
113

8.2 Introduction 

The oxysynthesis process is the hydroformylation of olefins with synthesis gas (CO and 

H2). It is the principal chemical route to C3-C15 aldehydes that are typically converted 

into alcohols, acids, or other derivatives. Worldwide production of oxo-aldehydes and 

alcohols was 6.5 million tons per year in 1997 (Bahrmann and Bach 2000). The 

announced planned capacity additions for oxo chemicals in the next 5 years would add 

almost an additional 1 million tons (Bitzzari et al. 2002). 

Hydroformylation is an industrial synthetic route for the conversion of a large number of 

olefins. It is used to produce solvents, synthetic detergents, flavorings, perfumes, other 

healthcare products, and other high value commodity chemicals. Oxo aldehyde products 

range from C3-C15. The higher end of this range (C11-C14) is often referred to as the 

detergent range because these aldehydes are typically used as intermediates in the 

synthesis of alcohols with the desired surfactancy and lack of crystallinity appropriate for 

detergents. 

The hydroformylation of olefins was discovered in 1938 by Roelen of Ruhrchemie as an 

outgrowth of Fischer-Tropsch Synthesis (FTS). Oxosynthesis conducted in Germany 

was thought to be a heterogeneously catalyzed reaction because of its association with 

FTS. Continuous addition of cobalt, however, was required for early oxosynthesis 

processes. This prompted additional investigation that identified soluble cobalt carbonyls 

as the active catalyst for the reaction making the oxosynthesis process one of the first 

significant homogeneously catalyzed industrial processes. Reactants, catalyst, and 

products are all present in the same phase, usually as liquids. Today, hydroformylation 

processes are the fourth largest commercial use of syngas (Wender 1996). 

Cobalt hydrocarbonyl was the first oxosynthesis catalyst commercially used. However, 

in the mid-1970s more active and selective ligand-modified rhodium-based catalysts were 

introduced for hydroformylation of lower olefins. For example, n-butanal from 

propylene is the highest volume oxo product. It can be converted to n-butanols (used as 

solvents) or 2-ethyhexanol, a plasticizer alcohol. 

8.3 Technology Description 

Oxosynthesis processes are very versatile and are used to commercially produce a wide 

range of products. The highest volume oxosynthesis process is the hydroformylation of 

propylene to butyraldehyde, which in turn, is converted to 2-ethylhexanol, a plasticizer 

alcohol used to make flexible PVC. Hydroformylation is a key step in the commercial 

production of Vitamin A. Hydroformylation of allyl alcohol is used to commercially 

produce butanediol. 
114

8.4 Chemistry 

The oxosynthesis process is also known as hydroformylation and involves the reaction of 

CO and H2 with olefinic hydrocarbons to form an isomeric mixture of normal- and isoaldehydes. The basic oxosynthesis reaction is highly exothermic and is 

thermodynamically favorable at ambient pressures and low temperatures (Whyman 

1985). The reaction proceeds only in the presence of homogeneous metal carbonyl 

catalysts. 

One of the more important factors in oxosynthesis is the normal-to-branched isomeric 

ratio (n/i). The normal, or straight chain isomer is the desired product. 

RCH=CH2 + CO + H2 → RCH2CH2CHO (normal) + R(CH3)CHCHO (branched) 

Usually a 1:1 H2:CO syngas mixture is required for oxosynthesis. The overall reaction 

rate has first-order dependence on the hydrogen partial pressure and inverse-first order 

dependence on CO partial pressure making the reaction rate essentially independent of 

total pressure (Pruett 1979). Higher CO partials pressures are usually required, however, 

to maintain the stability of the metal carbonyl catalysts. The reaction is also first order in 

olefin and metal concentration at the higher CO partial pressures. 

The mechanism for the homogeneously catalyzed hydroformylation reaction is a function 

of the catalyst system used in the process. However, the general steps in the reaction 

paths are similar. The first step is to remove CO from the catalyst organometallic 

complex to give a coordinatively electron deficient species. The olefin attaches to the 

vacated d-orbital at this site in the catalyst complex. The double bond in the olefin 

attaches to the metal atom yielding an alkyl metal carbonyl complex. The olefin inserts 

into the M-H bond. CO then inserts into the complex at the C-M bond followed by 

hydrogen insertion at the same point to yield the aldehyde product. A generalized 

reaction mechanism taken from Whyman (1985) follows: 

2 2 3 2 2 2 3 2

2 2 2 2 2

2 2 2 2

2 2 2

2 2

2 6 2 2 3 2

( ) ( ) ( )

( ) ( )

( ) ( )

( ) ( ) ( )

( ) ( ) ( )

( ) 2 ( ) 2 ( ) 2

RCH CH COM CO L HM CO L RCH CH CHO M CO L

RCH CH COM CO L H RCH CH CHO HM CO L

RCH CH M CO L CO RCH CH COM CO L

H RCH CH M CO L RCH CH M CO L

RCH CH HM CO L H RCH CH M CO L

M CO L H HM CO L HM CO L CO

x x

x x

x x

x x

x x

+ → + +

+ → +

+ ↔

= ↔

= + ↔ =

+ ↔ ↔ +

M= metal (Co or Rh) L = ligand (CO, PR3, or PPh3) x = 2 or 3 
115

8.5 Catalysts 

Oxosynthesis is a rapid reaction catalyzed by soluble cobalt or rhodium complexes. The 

advances made in organometallic chemistry in the 1950s and 1960s led to improved 

hydroformylation synthesis catalysts with higher thermal stability and greater selectivity. 

Three complimentary catalytic hydroformylation processes have been developed and 

commercialized. The choice of catalyst depends on the particular starting olefin or 

desired product. Phosphine-modified rhodium catalysts are best to produce 2-

ethylhexanol from propylene and syngas. Unmodified cobalt catalysts are versatile 

enough to be used for hydroformylation of high carbon number mixed olefins. 

Phosphine-modified cobalt catalysts are used for the production of higher, detergent 

range, alcohols rather than aldehydes. Each of these catalyst systems will be discussed. 

8.5.1 Cobalt Carbonyl 

The first hydroformylation catalysts were cobalt carbonyls, the specific active catalyst 

was found to be HCo(CO)4 in equilibrium with Co2(CO)8 based on the following 

reaction: 

Co2(CO)8 + H2 ↔ 2 HCo(CO)4

Cobalt metals and most cobalt salts will form cobalt carbonyl under hydroformylation 

conditions. The cobalt catalyzes both double bond isomerization and oxosynthesis. 

Undesired competing side reactions, such as the direct hydrogenation of the starting 

olefin and the condensation of product aldehydes to high boiling products, are generally 

avoided in the Co-catalyzed process. 

For cobalt carbonyl catalysts, a n/i ratio of 4:1 can be achieved with catalyst 

concentrations of 0.1-1% metal/olefin at 200-300 atm and 110-200°C with a 1:1 H2/CO 

ratio. Lower process temperatures and higher CO partial pressures favor the formation of 

the straight chain isomer, however, the overall conversion efficiency decreases. The 

cobalt carbonyl catalysts are also not very stable at the higher process temperatures. 

Catalyst tends to deposit on reactor walls decreasing catalyst activity and reducing 

catalyst recovery at elevated process temperatures. 

8.5.2 Phosphine-Modified Cobalt 

In the early 1960s Shell Oil Company commercialized a new cobalt-based 

hydroformylation process. The addition of a phosphine ligand to Co resulted in a 

trialkylphosphine-substituted cobalt carbonyl catalyst [HCo(CO)3P(n-C4H9)3]. This new 

catalyst had higher selectivity to straight-chain aldehydes (n/i = 7:1) with improved 

thermal stability compared to the unsubstituted cobalt catalysts. The improved thermal 

stability allows for lower process pressures but higher process temperatures. The Shell 

process conditions are 50-100 atm and 160-200°C with H2:CO = 1. Even though this 

catalyst has improved thermal stability, it has a lower hydroformylation activity than 
116

cobalt carbonyl catalysts, hence the higher reaction temperature. The higher 

temperatures also increase the competing olefin hydrogenation reaction. 

Shell has optimized this process to produce detergent range alcohols (C11-C14) in a single 

step by capitalizing on the conversion of terminal olefins to alcohols by hydrogenation of 

the aldehyde hydroformylation products. A high n/i ratio results from increased 

isomerization rates concurrently with hydroformylation with the phosphine-modified 

cobalt catalyst. 

8.5.3 Phosphine-Modified Rhodium 

Significantly lower operating pressures and temperatures and increased selectivity to 

linear products was demonstrated with the introduction of phosphine-modified rhodium 

catalysts. In 1976, Union Carbide and Davy Process Technology commercialized the LP 

Oxo process based on a triphenylphosphine-modified rhodium catalyst. With a 

HRh(CO)(PPh3)3 catalyst, the process is maintained at 7-25 atm and 60-120°C and a n/i 

ratio of 8-12:1 is possible. Low temperatures, high carbon monoxide partial pressure, 

and high ligand concentration on the Rh catalysts favors the formation of the straightchain isomers. Rh-based catalysts are mainly used for the hydroformylation of lower 

olefins (propylene to butyraldehyde, for example) but have limited use for higher olefins 

because of thermal instability of the catalyst at the high temperatures required for 

distillation to separate product and catalyst. Another important factor when selecting Rhbased hydroformylation catalysts is the high cost and low availability of rhodium. This 

makes catalyst separation and recycle an economically important process consideration. 

The high cost of rhodium, however, is offset by lower equipment costs, increased 

activity, and higher selectivity and efficiency. The development of water-soluble Rhbased catalysts avoids some of these issues. Rhone-Poulenc commercialized an oxo 

process based on a water-soluble Rh catalyst in 1984 (Billig and Bryant 2000). 

8.6 Gas Cleanliness Requirements 

Catalyst lifetimes are significantly reduced by poisoning from strong acids, HCN, 

organosulfur, H2S, COS, O2, and dienes (Bahrmann and Bach 2000). 

8.7 Status/Technology Developers 

The LP OxoTM Low Pressure Oxo Process is the result of a collaboration between Davy 

Process Technology (DPT) and Union Carbide Corporation that began in 1971. It is the 

world's leading process for the production of oxo alcohols from olefins. It represents 

most of the world's oxo alcohols capacity licensed during the last 20 years. A schematic 

of the LP Oxo Process is shown in Figure 10. 
117

Figure 10: Schematic Flow Diagram of the LP Oxo Process 

In the LP Oxo Process, propylene is reacted with synthesis gas to produce butyraldehydes 

(n/i = 10-20:1), which are in turn converted to 2-ethylhexanol. Normal or isobutyraldehydes are also converted to normal and iso-butanols and other derivatives. 

Propylene to mixed butyraldehyde efficiencies as high as 97.5% have been demonstrated 

at commercial scale. The efficiency is a function of feedstock purity with a higher purity 

feed resulting in a higher efficiency because of lower purge losses. 

The Shell Process is another commercial oxo synthesis process based on a phosphine -

modified cobalt catalyst to produce alcohols directly from olefins. The conversion of 

product oxo aldehydes to alcohols occurs in a single step in the Shell process because of 

the high hydrogenation activity of the modified catalyst. This process yields detergent 

range alcohols from a range of higher olefins. 

A summary of oxochemical producers in the U.S. as cited in the Chemical Economics 

Handbook (CEH) (Bitzzari et al. 2002) is presented in Table 27. 

Stirred-Tank

Reactor

Reprocessing

Purify

Purify

CO/H2

H2

Olefin

Rh complex

and ligand Offline Catalyst

Reprocessing

Compressor

Demister

Condenser

Separator

Stripping

Column

Product

Off gas
118

Table 27: Summary of U.S. Oxo Chemical Producers 

Company Plant Location Annual Capacity as of 9/01/02 (tonnes x103/yr) Oxo Products Catalyst Final Products BASF Corporation, Chemicals Division Industrial Organics 

Freeport, TX 

270 30 

n-butyraldehyde isobutyraldehyde 

Rh n-butanol 2-ethylhexano

neopentyl glycol; isobut

Celanese Ltd. Chemicals Division 

Bay City, TX 

45 250 27 

propionaldehyde n-butyraldehyde iosbutyraldehyde 

Rh 

n-propanol n-butanol isobutanol 

Dow Chemical Taft, LA 272 

15 

n-butyraldehyde isobutyraldehyde 

Rh n-butanol 

iso-butanol 

Dow Chemical Texas City, 

TX 

100 45 335 33 

propionaldehyde valeraldehyde n-butyraldehyde isobutyraldehyde 

Rh 

n-propanol; propionic acvaleric acid; amyl alcohn-butanol; 2-ethylhexanisobutanol 

Eastman Chemical Company, Eastman Division 

Longview, TX 

118 424 184 

propionaldehyde n-butyraldehyde isobutyraldehyde 

Rh 

n-propanol; propionic acn-butanol; 2-ethylhexanbutyric acid; methyl amyneopentyl glycol; methyacetate 

ExxonMobil Chemical Company 

Baton Rouge, LA 

408 Alcohols: 

C6-C13 branched 

C9-C11 linear 

Co phthalate and trimellitate 

Shell Chemical Company Geismar, 

LA 

340 C9-C15 linear alcohols Co-PR3 Neodol® alcohols 

Sterling Chemicals, Inc. Texas City, 

TX 

102 C7, C9, C11 linear 

alcohols Co Phthalate plasticizers 

Sunoco, Inc. Pasadena, 

TX 

140 12 

n-butyraldehyde isobutyraldehyde 

Rh 2-ethylhexanol 

sold without conversion 

Source: (Bitzzari et al. 2002) 
119

8.8 Cost 

The cost of oxo products are dependent on the initial costs of the olefin feedstocks, 

primarily propylene and ethylene. The oxo aldeyhdes are effectively intermediates for 

the production of alcohols, acids, and other chemical products. The market price for 

these products also varies with the cost of the feedstocks and intermediates. Chemical 

Economics Handbook (Bitzzari et al. 2002) contains estimated costs for a variety of oxo 

end products. The price of major oxochemical products listed above in Table 27 are 

presented in Table 28. Although not shown, the highest priced oxo chemical in the CEH 

list is pelargonic acid ($1.33/lb). The majority of the oxo products listed in CEH fall in 

the range of 50-70¢/lb. 

Table 28: U.S. Prices for Oxo Chemicals 

Product 2002 Price (¢/lb) 

n-butyraldehyde 52 

n-butanol 55 

2-ethylhexanol 63 

2-ethyl hexanoic acid 68 

isobutyraldehyde 52 

Isobutanol 50 

n-propanol 55 

propionic acid 44 

amyl alcohol 63 

Source: (Bitzzari et al. 2002) 

8.9 R&D Needs 

Continuing advances in improving the activity of hydroformylation catalysts are being 

made and should continue in the future. Thermal stability of future catalysts should be 

improved to continue the trend of reducing the oxosynthesis reaction severity. 

Optimizing the organometallic chemistry of these catalysts to increase activity and ease 

of catalyst separation from products is and area that would have a significant economic 

benefit to the overall oxosynthesis process. This is especially important for the Rh-based 

processes. Another area of interest is the development of heterogeneous 

hydroformylation catalysts, which would eliminate the challenges of catalyst separation 

and loss. 
120

8.10 References 

Bahrmann, H., and Bach, H. (2000). "Oxo Synthesis." Ullman's Encyclopedia of 

Industrial Chemistry, Wiley - VCH Verlag GmbH & Co. 

Billig, E., and Bryant, D. R. (2000). "Oxo Process." Kirk-Othmer Encyclopedia of 

Chemical Technology, John Wiley & Sons. 

Bitzzari, S. N., Gubler, R., and Kishi, A. (2002). "Oxo Chemicals." Chemical Economics 

Handbook, SRI International, 1-121. Report number 682.7000. 

Pruett, R. L. (1979). "Hydroformylation." Advances in Organometallic Chemistry, 17, 1-

60. 

Wender, I. (1996). "Reactions of synthesis gas." Fuel Processing Technology, 48(3), 189-

297. 

Whyman, R. (1985). "Industrial Applications of Homogeneous Catalysts." Selected 

Developments in Catalysis, J. R. Jennings, ed., Blackwell Scientific Publications, Oxford, 

128. 
121

9 Isosynthesis 

9.1 Summary 

The isosynthesis reaction converts syngas over a thorium or zirconium-based catalyst at 

relatively extreme (150-1000 atm and 450°C) conditions to isobutene and isobutane. 

Demand for MTBE as a gasoline additive increased in the early 1990s to the point that 

supply shortages of petrochemical isobutene were becoming a concern. This provided 

renewed interest in the conversion of syngas to i-C4 hydrocarbons, however, the 

isosynthesis process is not currently commercially practiced. Research and development 

efforts are being focused on finding isosynthesis catalysts that have high activity and high 

selectivity to isobutene at less severe process conditions, especially at lower pressure 

(~50 atm). 
122

Section Contents 

9.1 Summary ...................................................................................................................... 121

9.2 Introduction ................................................................................................................. 123

9.3 Technology Description............................................................................................... 123

9.4 Chemistry..................................................................................................................... 123

9.5 Catalysts...................................................................................................................... 124

9.6 Gas Cleanliness Requirements.................................................................................... 125

9.7 Status/Technology Developers .................................................................................... 125

9.8 R&D Needs.................................................................................................................. 125

9.9 References.................................................................................................................... 125
123

9.2 Introduction 

The isosynthesis reaction converts syngas over a thorium or zirconium-based catalyst at 

relatively extreme conditions to isobutene and isobutane. This reaction has often been 

considered a variation of Fischer-Tropsch Synthesis, however, there are major differences 

between the two synthetic processes (Feng et al. 1994). Isosynthesis is selective to i-C4

hydrocarbons whereas FTS forms a range of olefins according to the ASF distribution. 

Only trace amounts of oxygenates (water, methanol, isobutanol, DME, etc.) are formed 

under isosynthesis reaction conditions. Selective formation of branched hydrocarbons 

also occurs in isosynthesis. 

The isosynthesis reaction was first studied by Pilcher and Ziesecke in the early 1940s 

using ThO2-based catalysts and severe reaction conditions of 150-1000 atm and 450°C 

(Sofianos 1992). The goal was to optimize isobutane production for use in high-octane 

gasoline. The development of the petroleum industry led to waning interest in 

commercial isosynthesis processes and since these initial studies, isosynthesis has not 

received a lot of industrial attention. 

Demand for MTBE as a gasoline additive increased in the early 1990s to the point that 

supply shortages of petrochemical isobutene were becoming a concern (Sofianos 1992). 

This led to renewed interest in the isosynthesis reaction to convert syngas to branched 

hydrocarbons, particularly isobutene, which is a reactant in the synthesis of MTBE. 

Catalyst development has been focused on ZrO2-based catalysts for increased activity and 

selectivity to i-C4 products. 

9.3 Technology Description 

No commercial isosynthesis processes have been developed. Most laboratory studies 

have been conducted in gas-solid fixed bed reactors. Slurry reactors are also being 

investigated for isosynthesis, in fact, selectivities to C4 products are reported to be higher 

in slurry reactors compared to fixed bed reactors (Erkey et al. 1995). 

There is always a large amount of CO2 produced in the isosynthesis reaction, likely by 

the reaction of H2O and CO in the water-gas shift (Su et al. 2000a; Su et al. 2000b). CO2

recycle and reuse should be investigated to improve conversion efficiency to products. 

9.4 Chemistry 

Isosynthesis conditions are optimized for isobutene/isobutane production. At lower 

temperatures, alcohols and other oxygentates are formed. At higher temperatures, 

methane and aromatics are formed. 

The isosynthesis mechanism is related to that of higher alcohol synthesis. It involves two 

chain growth mechanisms — a step-wise CO insertion reaction and a condensation 

reaction mechanism involving surface adsorbed oxygenates. The precursor to the 
124

oxygenate products has been identified as a surface adsorbed methoxide species (Maruya 

et al. 1996). The oxygenates, alcohols and ethers, are more than likely the primary 

reaction products that undergo dehydration and hydrogenation to form the iso-alkenes 

and branched alkanes. These two competing chain growth mechanisms result in a 

discontinuity in the ASF distribution at C4 that explains the relatively high selectivity of 

the isosynthesis reaction to C4 products. 

9.5 Catalysts 

The first catalysts used by Pilcher and Ziesecke were ThO2-based. Unpromoted thoria 

catalysts have demonstrated 46% CO conversion at 150 atm and 450°C, but only 10% 

isobutene in the C4 fraction that is mostly isobutane (Sofianos 1992). DME was the main 

product at higher pressures. Thorium-based catalysts are good alcohol dehydration 

catalysts and, in turn, are the most active isosynthesis catalysts. They have long lifetimes 

because they can be regenerated by oxidizing the accumulated coke that deposits on the 

surface and, they are not poisoned by sulfur. Unfortunately, they are radioactive 

precluding their commercial use. 

Zirconium-based catalysts also have high activity for isosynthesis. Unpromoted 

zirconium catalysts have demonstrated 32% CO conversion at 150 atm and 450°C with 

much higher selectivity to isobutene (Sofianos 1992) compared to the thorium-based 

catalysts. The overall activity of ZrO2-based catalysts for isosynthesis is lower than 

ThO2-based catalysts. 

Various promoters have been investigated to improve the activity and selectivity of ThO2

and ZrO2 catalysts (Jackson and Ekerdt 1990a; Jackson and Ekerdt 1990b). The most 

active isosynthesis catalyst is 20% Al2O3/ThO2. Other promoters such as Zn, Cr and 

alkali metals have also been tested. The addition of alkali metals to zirconium catalysts 

had a negative effect on catalyst performance (Jackson and Ekerdt 1990b; Li et al. 

2001a). 

Doped zirconium-based catalysts posses oxygen vacancies in the oxide lattice. The most 

active catalysts tend to have maximum ionic conductivity suggesting that vacancies in the 

crystal lattice play an important role in the isosynthesis reaction (Li et al. 2001b; Li et al. 

2002). These oxygen vacancy sites are required for methoxide formation on the catalyst 

surface that contributes to the condensation reaction (Jackson and Ekerdt 1990a). 

The selectivity of the isosynthesis reaction depends on the nature of the active catalyst 

sites, including oxygen vacancies on the surface, and the number of acidic and basic sites. 

The balance between acidic and basic catalyst sites dictates overall activity and 

selectivity (Li et al. 2002). Enhancing the number of acidic sites on the ZrO2 catalyst 

increased activity and selectivity to linear C4 hydrocarbons. Increasing the number of 

basic sites on the catalyst increased the yield of iso-C4 hydrocarbons (Li et al. 2001b). 

The acidic catalyst sites are thought to promote condensation and dehydration reactions. 

The basic sites are known to catalyze CO insertion reactions. The activity of promoted 

isosynthesis catalyst systems is related to the acid/base ratio, which can be altered and 
125

controlled by varying the preparation procedure for mixed oxide catalysts (Feng et al. 

1995; Feng et al. 1994). 

9.6 Gas Cleanliness Requirements 

Thoria catalysts are not poisoned by sulfur and have high resistance to other poisons as 

well. 

9.7 Status/Technology Developers 

The isosynthesis process is not currently commercially practiced. 

9.8 R&D Needs 

Research and development efforts are focused on finding isosynthesis catalysts that have 

high activity and high selectivity to isobutylene at less severe process conditions, 

especially at lower pressure. Research catalysts are being developed for operation at 50 

atm compared to the original studies at 300 atm. 

9.9 References 

Erkey, C., Wang, J. H., Postula, W., Feng, Z. T., Philip, C. V., Akgerman, A., and 

Anthony, R. G. (1995). "Isobutylene Production from Synthesis Gas over Zirconia in a 

Slurry Reactor." Industrial & Engineering Chemistry Research, 34(4), 1021-1026. 

Feng, Z. T., Postula, W. S., Akgerman, A., and Anthony, R. G. (1995). "Characterization 

of Zirconia-Based Catalysts Prepared by Precipitation, Calcination, and Modified Sol-Gel 

Methods." Industrial & Engineering Chemistry Research, 34(1), 78-82. 

Feng, Z. T., Postula, W. S., Erkey, C., Philip, C. V., Akgerman, A., and Anthong, R. G. 

(1994). "Selective Formation of Isobutane and Isobutene from Synthesis Gas over 

Zirconia Catalysts Prepared by a Modified Sol-Gel Method." Journal of Catalysis, 

148(1), 84-90. 

Jackson, N. B., and Ekerdt, J. G. (1990a). "Isotope Studies of the Effect of Acid Sites on 

the Reactions of C-3 Intermediates During Isosynthesis over Zirconium Dioxide and 

Modified Zirconium Dioxide." Journal of Catalysis, 126(1), 46-56. 

Jackson, N. B., and Ekerdt, J. G. (1990b). "The Surface Characteristics Required for 

Isosynthesis over Zirconium Dioxide and Modified Zirconium Dioxide." Journal of 

Catalysis, 126(1), 31-45. 

Li, Y. W., He, D. H., Cheng, Z. X., Su, C. L., Li, J. R., and Zhu, Q. M. (2001a). "Effect 

of calcium salts on isosynthesis over ZrO2 catalysts." Journal of Molecular Catalysis aChemical, 175(1-2), 267-275. 
126

Li, Y. W., He, D. H., Yuan, Y. B., Cheng, Z. X., and Zhu, Q. M. (2001b). "Selective 

formation of isobutene from CO hydrogenation over zirconium dioxide based catalysts." 

Energy & Fuels, 15(6), 1434-1440. 

Li, Y. W., He, D. H., Yuan, Y. B., Cheng, Z. X., and Zhu, Q. M. (2002). "Influence of 

acidic and basic properties of ZrO2 based catalysts on isosynthesis." Fuel, 81(11-12), 

1611-1617. 

Maruya, K., Takasawa, A., Haraoka, T., Domen, K., and Onishi, T. (1996). "Role of 

methoxide species in isobutene formation from CO and H- 2 over oxide catalysts - 

Methoxide species in isobutene formation." Journal of Molecular Catalysis a-Chemical, 

112(1), 143-151. 

Sofianos, A. (1992). "Production of Branched-Chain Hydrocarbons Via Isosynthesis." 

Catalysis Today, 15(1), 149-175. 

Su, C. L., He, D. H., Li, J. R., Chen, Z. X., and Zhu, Q. M. (2000a). "Influences of 

preparation parameters on the structural and catalytic performance of zirconia in 

isosynthesis." Journal of Molecular Catalysis a-Chemical, 153(1-2), 139-146. 

Su, C. L., Li, J. R., He, D. H., Cheng, Z. X., and Zhu, Q. M. (2000b). "Synthesis of 

isobutene from synthesis gas over nanosize zirconia catalysts." Applied Catalysis aGeneral, 202(1), 81-89. 

 
127

10 Yields, Energy, and Economics 

10.1 Summary 

This section examines the yields, energy consumption, and economics for six of the 

products discussed in this report: hydrogen, methanol, fischer tropsch liquids, mixed 

alcohols, olefins and ethanol. The main focus of this section is on products from 

biomass. However, the yield data and energy consumption compares biomass to fossil 

based systems. The economics takes information from various biomass studies and puts 

it on a common basis so that a side-by-side economic comparison can be made. 

Overall, hydrogen has the highest yield from biomass followed by methanol. However, 

as expected, in general, the product yields from biomass resources are lower than if the 

product were produced from natural gas. The energy ratios were determined using a life 

cycle approach. Two energy ratios were calculated: a fossil energy ratio and a primary 

energy ratio. The fossil energy ratio is defined as (the energy in the fuel product)/(the 

fossil energy consumed by the system). The primary energy ratio includes the renewable 

energy consumption, and thus, is defined as (the energy in the fuel product)/(the fossil 

energy + the biomass energy consumed by the system). The fossil energy ratios for the 

biomass systems are considerably higher than those for the fossil systems (the fossil 

systems’ energy ratios are all less than one). Although biomass is renewable, when the 

biomass energy is included, the results show that the primary energy ratio is similar for 

the biomass and fossil systems. All biomass fuels have potential to significantly reduce 

the import of petroleum products. 

For all of the products examined in the economic analysis section (H2, MeOH, FTL, 

EtOH, mixed alcohols, olefins), syngas production accounts for at least 50% of the 

product cost and in many cases it is more like 75%. The economics show hydrogen and 

methanol to be the most cost competitive products from biomass. Ethanol from biomassderived syngas could potentially be cost competitive. However, the state of this 

technology is still at the lab scale and the amount of data available is limited. Therefore, 

the ethanol analysis has a higher level of uncertainty than the other products examined. 

Additionally, because of the limited amount of design data, more analysis should be 

performed for mixed alcohols synthesis to examine biomass-optimized configurations 

including recycle for maximum conversion and the resulting economics. 
128

Section Contents 

10.1 Summary .................................................................................................................... 127

10.2 Product Yields ........................................................................................................... 129

10.3 Life Cycle Energy Ratios........................................................................................... 131

10.4 Petroleum Product Replacement............................................................................... 133

10.5 Economics.................................................................................................................. 135

10.6 References.................................................................................................................. 141

List of Figures 

Figure 11: Product Yields from Fossil and Biomass Resources ......................................... 130 

Figure 12: Life Cycle Fossil Energy Ratio.......................................................................... 132 

Figure 13: Life Cycle Primary Energy Ratio ...................................................................... 132 

Figure 14: Potential Replacement of Petroleum Products with Biomass Fuels (Energy 

Basis)................................................................................................................. 134 

Figure 15: Product Price in $/GJ (LHV)............................................................................. 139 

Figure 16: Capital Cost Breakdown for Biomass to MeOH Systems.................................. 140

List of Tables 

Table 29: Feedstock to Product Material Balance Information.......................................... 129 

Table 30: Life Cycle Energy Ratios (LHV basis)................................................................. 131 

Table 31: Economic Parameters for Discounted Cash Flow Analysis................................ 136 

Table 32: Summary of Cases Examined .............................................................................. 137 

Table 33: Base Case Minimum Product Selling Price ........................................................ 138 

Table 34: Price of Commercial Products............................................................................ 138 

Table 35: Sensitivity – Minimum Product Selling Price...................................................... 139
129

10.2 Product Yields 

Product yields vary, depending on the process configuration (e.g. once through or 

recycling to maximize yields), technology used (e.g., gas versus liquid phase), and type 

of feedstock. Table 29 gives the product yields from both fossil and biomass resources 

along with the references used to calculate the yields. In some instances, material 

balance data from NREL’s gasifier models (Spath, et al, 2000) were used in conjunction 

with material balance information from fossil technologies in order to come up with a 

biomass-to-fuels process. To compare the magnitude of the numbers, Figure 11: Product 

Yields from Fossil and Biomass Resources, is a graph of the GJ of product/GJ of feed 

(column 4 of Table 29). As expected, generally, the yield of any product from biomass is 

lower than that product derived from natural gas. Overall, hydrogen has the highest yield 

from biomass followed by methanol while mixed alcohols have the lowest product yield 

from biomass. It should be noted, however, that the there was a limited amount of data 

from which to derive the material balance for the ethanol and mixed alcohols products 

from biomass and that these processes are not optimized. 

Table 29: Feedstock to Product Material Balance Information 

System 

(BD = bone dry) 

kg of 

product/Mg 

of BD feed 

GJ of 

product/Mg 

of BD feed 

GJ of 

product/GJ of 

feed (LHV) 

Reference 

Natural gas to H2 337 40 0.83 Leiby, 1994 

Coal to H2 96 12 0.44 Spath and Amos (2000) 

Biomass to H2 52-105 6-13 0.36-0.73 Spath, et al. (2000); 

Hamelinck, et al. (2001) 

Natural gas to MeOH 1,445 29 0.61 SRI (2002) 

Coal to MeOH 518-666 11-14 0.40-0.52 Bailey (1979); Michaels 

(1979); SRI (2002) 

Biomass to MeOH 246-554 5-11 0.29-0.65 Hamelinck, et al. (2001) 

Natural gas to FT 570-666 26-31 0.54-0.63 Navqvi (2000) 

Biomass to FT 60-159 3-7 0.16-0.43 Tijmensen, et al. (2000); 

Hamelinck, et al, 2003; NREL 

gasification material balance 

info (Spath, et al, 2000) with 

Apanel (Dec. 2002) 

Natural gas to mixed 

alcohols 

1,210 31 0.63 Nirula (1994) 

Biomass to mixed 

alcohols 

135-223 3-6 0.18-0.30 NREL gasification material 

balance info (Spath, et al, 

2000) with Nirula (1994) 

Natural gas to olefins 453 21 0.43 Apanel (2002) 

Biomass to MeOH then 

olefins 

77-174 4-8 0.21-0.46 MeOH material balances with 

Apanel (2002) 

Biomass to syngas then 

ethanol 

228 7 0.35 Putsche (1999) 
130

Figure 11: Product Yields from Fossil and Biomass Resources 
131

10.3 Life Cycle Energy Ratios 

The energy balance of a system can be examined in a number of ways. This study looked 

at the energy balance of each system using a life cycle approach. Two energy ratios were 

calculated: 1) the fossil energy ratio which is defined as the energy in the product 

divided by the total fossil energy consumed from feedstock procurement through product 

generation and 2) the primary energy ratio which includes the energy in the biomass 

feedstock even though it is renewable. Table 30 gives the fossil and primary energy 

ratios for several products. Most of these ratios (except as noted) were derived using the 

material balance information in Table 29 along with the life cycle information contained 

in Mann and Spath (1997) for producing and transporting a dedicated biomass feedstock, 

specifically poplar. Figure 12 and Figure 13 are graphical representations of the 

numbers in Table 30. Note that Figure 12 is a log scale so that the lifecycle energy ratio 

for the fossil systems, which are all less than one, could be seen. Obviously, the energy 

ratios will follow the same general pattern as the product yields given in the previous 

section (e.g., hydrogen from biomass has the highest energy ratio followed by methanol). 

The fossil energy ratios for the biomass systems are considerably higher than those for 

the fossil systems because in the fossil systems, the main source of energy consumption 

is the feedstock itself. Thus, there is a considerable savings in fossil energy consumption 

by using fuels from biomass-derived syngas. Although biomass is a renewable resource, 

when the energy content of the biomass is included in the energy ratio, the magnitude of 

the number is similar to that for the fossil based system. 

Table 30: Life Cycle Energy Ratios (LHV basis) 

System Fossil Energy Ratio 

(Efuel product/Efossil energy 

consumed) 

Primary Energy Ratio 

(Efuel product/(Efossil + E biomass energy 

consumed)) 

Natural gas to H2 (a) 0.66 0.66 

Coal to H2 0.39 0.39 

Biomass to H2 14-29 0.35-0.71 

Natural gas to methanol 0.44 0.44 

Coal to methanol 0.34-0.44 0.34-0.44 

Biomass to methanol 12-26 0.29-0.64 

Natural gas to FT 0.39-0.46 0.39-0.46 

Biomass to FT 6-17 0.16-0.42 

Biomass to ethanol via syngas 16 0.35 

Natural gas to mixed alcohols 0.46 0.46 

Biomass to mixed alcohols 8-13 0.18-0.29 

Natural gas to olefins 0.31 0.31 

Biomass to methanol then olefins 8-19 0.20-0.45 

Gasoline/diesel (b) 0.83 0.83 

Source: (a) Spath and Mann (2000); 

(b) General Motors Corporation, Argonne National Laboratory. (2001) 
132

Figure 12: Life Cycle Fossil Energy Ratio 

Figure 13: Life Cycle Primary Energy Ratio 

Fossil energy ratio (Eproduct/Efossil)

Cross hatched area indicates a range of numbers

Light green indicates higher level of uncertainty

biomass

(via

syngas)

0.1

1

10

100

nat gas coal biomass nat gas coal biomass nat gas biomass nat gas biomass nat gas biomass corn

Hydrogen MeOH FTL Mixed Olefins EtOH

Alcohols

Note: Log scale Higher level of 

uncertainty

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

nat gas coal biomass nat gas coal biomass nat gas biomass nat gas biomass nat gas biomass biomass

(via

syngas)

Primary energy ratio (Eproduct/(Efossil +Ebiomass))

Cross hatched area indicates a range of numbers

Light green indicates higher level of uncertainty

corn

Hydrogen MeOH FTL Olefins EtOH

Mixed 

Alcohols

Higher level of 

uncertainty
133

10.4 Petroleum Product Replacement 

Using the material balance information in Table 29 (specifically, the highest yield for 

each product), calculations were made to examine the potential petroleum product 

replacement from biomass-derived fuels. Calculations were made using the current 

amount of biomass available in the U.S. that is less than $50/dry ton delivered. This 

amount of biomass is predicted to be 512 million tons/yr, from Oakridge National 

Laboratory, (http://bioenergy.esd.ornl.gov/papers/misc/resource_estimates.html). It is 

estimated that this could potentially be increased to 800 million tons/yr in the next two 

decades. Note that each product shown in Figure 14 assumes that all of the biomass (512 

million tons/yr) is used to produce that one product (i.e., 512 million tons converted to H2

or 512 million tons converted to MeOH). 

The columns shown for each product in Figure 14 are: 

• % of total petroleum products consumed in 2000 that the biomass-derived product could 

replace 

• % of yearly net petroleum imports in 2000 that the biomass-derived product could 

account 

• % of the increase in petroleum products consumed in 2020 compared to 2000 that the 

biomass-derived product could replace 

• % of the increase in yearly net imports in 2020 compared to 2000 that the biomassderived product could account for. 

For example, the figure shows that if all of the biomass was converted to hydrogen, this is 

equivalent to 17% of the petroleum products consumed in 2000 or 30% of the net imports 

in that year. Additionally, this amount of hydrogen is equivalent to 47% of the predicted 

increase in petroleum products in 2020 compared to 2000. This also means that this 

quantity of hydrogen could account for 44% of the increase in predicted net imports in 

2020 compared to 2000. 

Figure 14 also shows that hydrogen has the greatest potential to replace petroleum 

products followed by methanol, with mixed alcohols being the smallest. Note however, 

that the comparison was made on an energy equivalent basis and does not take into 

account vehicle energy efficiency or fuel economy. In some cases, this could be a 

significant advantage. For example, although highly debated, hydrogen fuel cell vehicles 

are expected to be two to three times as energy efficient as current combustion engine 

vehicles. Thus, on a per vehicle mile traveled basis, hydrogen from biomass would 

replace a greater amount of petroleum products than the amount shown in Figure 14. 
134

Figure 14: Potential Replacement of Petroleum Products with Biomass Fuels (Energy Ba

0%5%

10%15%20%25%30%35%40%45%50%

hydrogen methanol olefins FTL ethanol - syngas

then fermentation

mix

% of yearly total petroleum products consumed in 2000% of yearly net petroleum imports in 2000% of increase (2020-2000) in predicted yearly total petrconsumed% of increase (2020-2000) in predicted yearly net petro

(Assumes 512 MM tons/yr of available biomass which goes to 
135 

10.5 Economics 

Many other groups have done economic analyses for the various products discussed in this 

report. However, these analyses are not all on the same basis. They use different plant sizes, 

different feedstock costs, varying economic parameters, and different methods for 

determining product price. As a result, a side-by-side economic comparison of the different 

products is difficult. In this study, an attempt was made to examine several products using 

the same approach with consistent economic parameters. 

In general, when enough detailed information existed, material and energy balances, along 

with capital and operating cost data, were extracted from studies by others. This information 

was input into a discounted cash flow rate of return spreadsheet in which the same plant sizes 

and economic parameters were used to determine the minimum product selling price required 

to obtain a specific after-tax rate of return. In some instances, energy and material balance 

data from NREL’s gasifier models (Spath, et al, 2000) were used in conjunction with energy 

and material balance information from fossil technologies to come up with a biomass-to-fuels 

process. The specific parameters used in this analysis are given in Table 31. In general, 

these are the same parameters as those used in NREL’s design report “Lignocellulosic 

Biomass to Ethanol Process Design and Economic Utilizing Co-Current Dilute Acid 

Prehydrolysis and Enzymatic Hydrolysis for Corn Stover” (Aden, et al., 2002). Additionally, 

the gasifier costs were kept consistent from product to product, using the costs for the 

Battelle/Ferco low-pressure indirectly heated gasifier and for the Institute of Gas Technology 

high-pressure gasifier used in NREL’s previous analyses (Spath, et al. 2000). 

Spreadsheets were set up for 5 hydrogen cases, 4 methanol cases, 8 fischer tropsch cases, 1 

ethanol case, 2 mixed alcohol cases, and 4 olefin cases. Table 32 briefly outlines the process 

configurations examined using the references given in Table 29. 

For the base case assumptions outlined in Table 31, the resulting minimal selling price for 

hydrogen, methanol, FTL, ethanol, mixed higher alcohols, and olefins via biomass derived 

syngas are shown in Table 33. 
136

Table 31: Economic Parameters for Discounted Cash Flow Analysis 

Parameter Base Case Sensitivity 

Values 

Plant life 20 years --- 

Plant size/biomass feed rate 91.9 BD ton/hr 23 & 60 BD ton/hr 

Discount rate 10% 15% & 20% 

Financing 100% equity 50% & 75% equity 

Plant depreciation 200% DB --- 

Plant recovery period 7 years --- 

Federal tax rate 39% --- 

Construction period 2.5 years --- 

Start-up time 6 months --- 

Feedstock cost $30/dry ton $15 & $42/dry ton 

On stream time 8380 hours --- 

Loan interest 7.5% --- 

Loan term 10 years --- 

Reference year 2001 --- 

Electricity price 4.1¢/kWh --- 

Direct costs (including equip installation, 

instruments and controls, piping, electrical and 

misc buildings) 

61% of total purchased 

equipment costs 

(TPEC) 

--- 

Total installed equipment cost (TIEC) TPEC + direct costs --- 

Warehouse 

Site development 

1.5% of TIEC 

4.5% of TIEC 

--- 

Total installed cost (TIC) TIEC + warehouse + 

site development 

--- 

Indirect Field Costs 

- field expenses 

- home office & construction fee 

- project contingency 

- prorateable costs 

 

20% of TIC 

25% of TIC 

3% of TIC 

10% of TIC 

--- 

Total capital investment (TCI) TIC + indirect costs --- 

Other costs (startup, permits, etc.) 10% of TCI --- 

Total Project Investment (TPI) TCI + other costs --- 

Note: This is the same basis as that used in Aden, et al., (2002). 
137 

Table 32: Summary of Cases Examined 

Product Cases 

H2 - IGT, HGCU, shift, PSA, combined cycle 

- BCL, wet gas cleaning, stm reformer, shift, PSA 

- BCL, wet gas cleaning but no tar cracker, stm reformer, shift, PSA 

- BCL, wet gas cleaning, shift, PSA, combined cycle 

- IGT, HGCU, reforming, shift, PSA 

MeOH - IGT, HGCU, ATR, LPMeOH, combined cycle 

- BCL, wet gas cleaning, stm reformer, LPMeOH, stm cycle 

- IGT, HGCU, ATR, shift, conventional MeOH, stm turbine 

- BCL, wet gas cleaning, stm reformer, shift, conv. MeOH, stm turbine 

FTL - BCL, wet gas cleaning, ATR, CO2 removal, slurry phase FT, hydrocracker, 

distillation 

- IGT, HGCU, ATR, CO2 removal, slurry phase FT, hydrocracker, distillation 

- IGT, wet gas cleaning, shift, slurry phase FT, product recovery and upgrading via 

distillation and hydrocracking, combined cycle (once thru) 

- IGT, wet gas cleaning, shift, solid bed FT, product recovery and upgrading via 

distillation and hydrocracking, combined cycle (once thru) (same process @70% 

conversion and 3 different alphas; which is the probability of chain growth) 

- IGT, wet gas cleaning, solid bed FT reactor, ATR & shift on recycle stream, 

product recovery and upgrading via distillation and hydrocracking, combined cycle 

- IGT, wet gas cleaning, solid bed FT reactor, shift on recycle stream, product 

recovery and upgrading via distillation and hydrocracking, combined cycle 

Ethanol via syngas 

fermentation 

- BCL, wet gas cleaning, fermentation, distillation, etoh dehydration 

Mixed higher alcohols - BCL, wet gas cleaning, CO2 separation, alcohol synthesis, product purification, 

CO2 separation, CO recovery 

- IGT, HGCU, CO2 separation, alcohol synthesis, product purification, CO2

separation, CO recovery 

Olefins (propylene) - MeOH via the 4 schemes listed above followed by MeOH conversion and 

propylene recovery via distillation 

Notes: 

(1) HGCU = granular beds, candle filter, alkali removal, halogen absorption, sulfur absorption 

(2) wet gas cleaning = tar cracker, cyclone, gas cooling, bag filter, COS hydrolysis, scrubbers, 

ZnO bed 
138

Table 33: Base Case Minimum Product Selling Price 

Product $/GJ 

(HHV) 

$/GJ 

(LHV) 

$/gallon or 

$/kg 

Reference where data 

was obtained 

H2 8-14 9-17 1.1-2.0/kg Spath, et al, 2000 

Hamelinck and Faaji, 2001 

MeOH 12-13 13-14 0.80-0.91/gal Hamelinck and Faaji, 2001 

FTL 19-25 20-27 2.7-3.6/gal Tijmensen, et al, 2000 

Hamelinck, et al, 2003 

NREL gasification material 

balance info (Spath, et al, 

2000) with Apanel (Dec. 2002) 

Ethanol via syngas 

fermentation 

14 16 1.3/gal Putsche, 1999 

Mixed higher alcohols 27-31 31-36 2.4-2.8/gal NREL gasification material 

balance info (Spath, et al, 

2000) with Nirula (1994) 

Olefins (propylene) 17-19 18-20 0.81-0.92/kg MeOH material balances with 

Apanel (2002) 

Table 34 presents the price of many of today’s commercial products. Figure 15 plots the 

base case ranges (from Table 33) and the commercial product prices in terms of $/GJ, LHV 

basis. Although, in general, the biomass-derived product prices are higher than the 

commercial prices, several of the products appear to be cost competitive: hydrogen, 

methanol, and ethanol. However, the ethanol analysis was based on a limited amount of 

data, so the level of uncertainty is higher than it is for the other products examined, thus the 

results should be viewed with caution. 

Table 34: Price of Commercial Products 

System $/GJ 

(HHV) 

$/GJ 

(LHV) 

$/gallon 

(except for H2) 

Comments 

Natural gas 3.4 3.8 N/A Average Nov. 2002 price 

Gasoline 5.6 6.1 0.74 

Diesel 5.1 5.5 0.75 

Excluding distribution, marketing, and 

taxes (Nov. 2002) 

Hydrogen 5-10 6-12 0.7-1.4/kg Plant gate price of SMR 

Methanol 4-10 5-12 0.3-0.7 General range of spot prices from 

1993-2001 

Ethanol 11-16 12-17 1.00-1.40 Price range from 1989-2000 

Ethanol Excise 

Tax Exemption 

6 7 0.53 Currently expected to expire in 2007 

Propylene 6-10 6-11 0.29-0.50/kg Apanel 2002 and website listed 

below*

*http://www.the-innovation-group.com/ChemProfiles/Propylene.htm
139 

Figure 15: Product Price in $/GJ (LHV) 

Sensitivity runs were performed by making changes to the base case values regarding plant 

size, feedstock cost, internal rate of return, and financing (debt/equity). A combination of 

options were examined. For each plant size, the spreadsheet was evaluated 27 times (using 

macros) to account for the combination of changes in the feedstock cost ($15, $30, and 

$42/dry ton), internal rate of return (10%, 15%, and 20%), and financing (50%, 75%, and 

100% equity). Table 35 gives the range of results for each product and all of the cases 

examined. Smaller plant sizes, higher rates of return, and higher feedstock costs will 

contribute to a higher product price compared to the base case. Conversely, a higher amount 

of debt and lower feedstock costs will result in a product price that is lower than the base 

case. 

Table 35: Sensitivity – Minimum Product Selling Price 

Product $/GJ 

(HHV) 

$/GJ 

(LHV) 

$/gallon or 

$/kg 

H2 6-42 7-48 0.84-5.9/kg 

MeOH 9-47 10-51 0.62-3.2/gal 

FTL 15-63 16-69 2.2-9.3/gal 

Ethanol via syngas 

fermentation 

12-32 13-35 1.0-2.8/gal 

Mixed higher alcohols 22-80 25-30 2.2-9.3/gal 

Olefins (propylene) 13-70 13-75 1.9-7.1/kg 

0

5

10

15

20

25

30

35

40

from

biomass

H2 from

biomass

MeOH from

biomass

diesel/ 

gasoline

from

biomass

(30% MeOH)

from

biomass

propylene from

biomass

(via

syngas)

corn

Product Price ($/GJ, LHV)

biomass derived product price

commercial product price

Higher level of 

uncertainty

Hydrogen MeOH FTL

Olefins EtOH

Mixed 

Alcohols

(excl. dist., 

marketing, 

& taxes)

(excl. 

excise 

tax 

credit)
140

Figure 16: Capital Cost Breakdown for Biomass to MeOH Systems 

For all of the products examined in the economic analysis (H2, MeOH, FTL, EtOH, mixed 

alcohols, olefins), syngas production accounts for at least 50% of the product cost and in 

many cases, it is almost 75%. Figure 16 shows a breakdown of the capital cost for methanol 

production from biomass for both a direct and an indirect gasification system (Taken from 

costs in Hamelinck and Faaij, 2001). Clean conditioned syngas (a combination of feed 

preparation, gasification, oxygen production, gas clean up and cooling, syngas compression 

and reforming) accounts for 60-64% of the total capital. Overall, steps should be made to 

optimize the biomass-to-fuels process in order to obtain the highest yield, least cost 

configuration. To reduce costs, efforts should be focused on minimizing the cost of clean 

syngas production. Economies of scale can play a large factor in lowering the product cost, 

therefore, opportunities to co-feed with coal or natural gas systems may be one way to get 

renewable fuels into the marketplace, just as co-firing biomass with coal is being done in the 

power generation industry. 

Direct system: GTI, HGCU, ATR, shift, conv MeOH, 

stm turbine

Indirect system: BCL, scrubbing, stm reformer, shift, 

conv MeOH, stm turbine

Clean conditioned syngas generation = 

60-64% of total capital

feed prep

gasification

O2 plant (Case 1 only)

clean up & cooling

syngas compression (Case 2 only)

reforming

CO2 removal

MeOH synthesis & compression

power generation

7%

14%

15%

9%

15%

9%

22%

9%

8%

6%

19%

9%

22%

5%

27%

4%
141 

10.6 References 

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

L.; Slayton, A.; Lucas, J. (2002). “Lignocellulosic Biomass to Ethanol Process Design and 

Economics Utilizing Co-Current Dilute Acid Prehydrolysis and Enzymatic Hydrolysis for 

Corn Stover.” National Renewable Energy Laboratory report NREL/TP-510-32438. 

Apanel, G.; Netzer, D. (2002). “Methanol to Propylene by the Lurgi MTP Process”. PEP 

Review 98-13, SRI International, Menol Park, CA. 

Apanel, G.J. (Dec. 2002). “Gas to Liquids Update”. PEP Review, Report No. 247, SRI 

International, Menol Park, CA. 

Bailey, E. (1979). "The coal to methanol alternative." Coal Technol. (Houston), 2nd(3), 149-

61. 

Fuel Ethanol Terminal Market Price - 18-month History. OXY-FUEL News Price Report. 

1995-2003 Hart Publications. 

Gasoline and Diesel Fuel Update (http://tonto.eia.doe.gov/oog/info/gdu/gasdiesel.asp ) 

General Motors Corporation, Argonne National Laboratory. (2001). “Well-to-Wheel Energy 

Use and Greenhouse Gas Emissions of Advanced Fuel/Vehicle Systems” 

Gray, D. and Tomlinson, G. (1997). "A technical and economic comparison of natural gas 

and coal feedstocks for Fischer-Tropsch synthesis." Studies in Surface Science and Catalysis

107(Natural Gas Conversion IV): 145-150. 

Hamelinck, C. N., and Faaij, A. P. C. (2001). "Future prospects for production of methanol 

and hydrogen from biomas." NWS-E-2001-49; ISBN 90-73958-84-9, Utrecht University. 

Hamelinck, C. N.; Faaij, A. P. C.; Uil, H,; Boerrigter, H. (2003). "Production of FT 

transportation fuels from biomass; technical options, process analysi and optimization, and 

development potential." NWS-E-2003-08; ISBN 90-393-3342-4, Utrecht University. 

Leiby, S.M. (1994). Options for Refinery Hydrogen. SRI Report No. 212. Menlo Park, CA. 

Mann, M.K.; Spath, P.L. (1997). Life Cycle Assessment of a Biomass Gasification 

Combined-Cycle Power System. National Renewable Energy Laboratory, Golden, CO, TP430-23076. 

Mann, M.; Spath, P. (2002). “Life Cycle Assessment Comparisons of Electricity From 

Biomass, Coal, and Natural Gas.” AICHE 2002 Annual Meeting, Indianapolis, IN, 

November 3-8. 

Michaels, H. J. (1979). "Methanol from coal using the Koppers-Totzek process." Coal 

Technol. (Houston), 2nd(3), 163-78. 
142

Navqvi, S.N. (2000). “Opportunities for Gas-to-Liquid Technologies” Supplement C. 

Process Economics Program. Menlo Park, California. 

Nirula, S. C. (1994). "Dow/Union Carbide Process for Mixed Alcohols from Syngas." PEP 

Review no. 85-1-4, SRI International, Menol Park, CA. 

Putsche, V. (1999). “Complete Process and Economic Model of Syngas Fermentation to 

Ethanol.” NREL Milestone Report. 

Spath, P.L; Mann, M.K. (November 2000). Life Cycle Assessment of Hydrogen Production 

via Natural Gas Steam Reforming. National Renewable Energy Laboratory, Golden, CO, 

TP-570-27637. 

Spath, P.L.; Lane, J.M; Mann, M.K; Amos, W.A. (April 2000). Update of Hydrogen from 

Biomass - Determination of the Delivered Cost of Hydrogen. NREL Milestone report. 

Spath, P.L.; Amos, W.A. (2000). “Hydrogen Production from Western Coal Including CO2

Sequestration and Coalbed Methane Recovery: Economics, CO2 Emissions, and Energy 

Balance”. Advances in Hydrogen Energy. New York: Plenum Publishers, pp. 17-30. 

SRI International. (1994). Options for Refinery Hydrogen. Report No. 212. Menlo Park, 

CA. 

Suresh, B.; Gubler, R.; Sasano, T. (2001). "Hydrogen." Chemical Economics Handbook 

Product Review, SRI International, Menlo Park, CA. Reprot number 743.5000. 

Tijmensen, M. J. A., Faaij, A. P. C., Hamelinck, C. N. and van Hardeveld, M. R. M. (2000) 

"The production of Fischer Tropsch liquids and power through biomass gasification", M. Sc. 

Thesis, Utrecht University/Science, Technology and Society, Utrecht. Report No. NWS-E2000-29. 

U.S. Natural Gas Prices and Natural Gas Summary from the Short-Term Energy Outlook 

(http://www.eia.doe.gov/oil_gas/natural_gas/info_glance/prices.html ) 

Wham, R. M., and Forrester III, R. C. "Available Technology for Indirect Conversion of 

Coal to Methanol and Gasoline: A Technology and Economics Assessment." Alternative 

energy sources III, Miami Beach, FL, 3-18. 
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1. AGENCY USE ONLY (Leave blank) 2. REPORT DATE

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Technical Report

4. TITLE AND SUBTITLE

Preliminary Screening – Technical and Economic Assessment of Synthesis Gas to 

Fuels and Chemicals with Emphasis on the Potential for Biomass-Derived Syngas 

6. AUTHOR(S) 

P.L. Spath and D.C. Dayton 

5. FUNDING NUMBERS

BBB3.4210 

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National Renewable Energy Laboratory 

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13. ABSTRACT (Maximum 200 words)

In principle, syngas (primarily consisting of CO and H2) can be produced from any hydrocarbon feedstock, including: natural gas, naphtha, 

residual oil, petroleum coke, coal, and biomass. The lowest cost routes for syngas production, however, are based on natural gas, the 

cheapest option being remote or stranded reserves. Economic considerations dictate that the current production of liquid fuels from syngas 

translates into the use of natural gas as the hydrocarbon source. Nevertheless, the syngas production operation in a gas-to-liquids plant 

amounts to greater than half of the capital cost of the plant. The choice of technology for syngas production also depends on the scale of 

the synthesis operation. Syngas production from solid fuels can require an even greater capital investment with the addition of feedstock 

handling and more complex syngas purification operations. The greatest impact on improving the economics of gas-to liquids plants is 

through 1) decreasing capital costs associated with syngas production and 2) improving the thermal efficiency with better heat integration 

and utilization. Improved thermal efficiency can be obtained by combining the gas-to-liquids plant with a power generation plant to take 

advantage of the availability of low-pressure steam. 

The extensive research and development efforts devoted to syngas conversion to fuels and chemicals are documented in a vast amount of 

literature that tracks the scientific and technological advancements in syngas chemistry. The purpose of this report is to review the many 

syngas to products processes and summarize the salient points regarding the technology status and description, chemistry, catalysts, 

reactors, gas cleanliness requirements, process and environmental performances, and economics. Table 1 lists the products examined in 

this study and gives some facts about the technology as well as advantages and disadvantages. Table 2 summarizes the catalysts, process 

conditions, conversions, and selectivities for the various syngas to products processes. Table 3 presents catalyst poisons for the various 

products. 

14. SUBJECT TERMS 15. NUMBER OF PAGES

biofuels; ethanol; fuels; chemicals; synthesis gas; syngas; economic assessment; 

technical assessment; technoeconomic assessment; biomass 16. PRICE CODE

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