Reactions Of Synthesis Gas

Fuel Processing Technology 48 (I 996) 189-297

Reactions of synthesis

gas

Irving Wender Chemicd

und Prtrolrum

Engineering

Depurtment,

University

oj’Pitt.shurgh.

Pittsburgh,

PA 15261.

USA

Abstract The use of synthesis gas (syngas) offers the opportunity to furnish a broad range of environmentally clean fuels and chemicals. There has been steady growth in the traditional uses of syngas. Almost all hydrogen gas is manufactured from syngas and there has been a tremendous spurt in the demand for this basic chemical; indeed, the chief use of syngas is in the manufacture of hydrogen for a growing number of purposes. Methanol not only remains the second largest consumer of syngas but has shown remarkable growth as part of the methyl ethers used as octane enhancers in automotive fuels. The Fischer-Tropsch synthesis remains the third largest consumer of syngas, mostly for transportation fuels but also as a growing feedstock source for the manufacture of chemicals, including polymers. Future growth in Fischer-Tropsch synthesis may take place outside the continental United States. The hydroformylation of olefins (the 0x0 reaction), a completely chemical use of syngas, is the fourth largest use of carbon monoxide and hydrogen mixtures; research and industrial application in this field continue to grow steadily. A direct application of syngas as fuel (and eventually also for chemicals) that promises to increase is its use for Integrated Gasification Combined Cycle (IGCC) units for the generation of electricity (and also chemicals) from coal, petroleum coke or heavy residuals. In the period 2005-2015, the amount of syngas employed in this manner may approach that used for all other specific purposes. Syngas is the principal source of carbon monoxide, which is used in an expanding list of so-called carbonylation reactions.

0. Overview

Synthesis gas (syngas), a mixture of hydrogen and carbon monoxide, can be manufactured from natural gas, coal, petroleum, biomass and even from organic wastes, so that sources of syngas are ubiquitous in nature. The availability and flexibility of the resource base are keys to the present and future uses of syngas and of its separate components, hydrogen and carbon monoxide. Syngas is a present and increasing source of environmentally clean fuels and chemicals and is also a potentially major fuel for the 037%3820/96/$15.00 PII

SO378-3820(96)01048-X

Copyright 0 1996 Elsevier Science B.V. All rights reserved.

190

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Technology 48 (1996) 189-297

production of essentially pollution-free electricity, as sulfur and nitrogen in parts per million can be removed from syngas relatively easily. This report will list the commercial uses of syngas in the order of their importance, compare them with uses of a decade ago and outline what promises to be a bright future for this mixture of gases. There are literally thousands of published papers and proceedings of various conferences that deal with syngas reactions. A very large number of these publications deals with catalytic reactions of syngas in an exploratory way aimed often at the elucidation of mechanistic pathways. These researches provide an understanding of various aspects of syngas reactions and point the way to possible future commercial applications and to new uses of syngas; they will be referred to when they are of particular pertinence. References to some of the large number of patents that deal with the chemistry and uses of syngas will be referred to occasionally but are not a major source of material for this review. The birth of syngas chemistry occurred in the early part of the 20th century. Methane was produced by the hydrogenation of carbon monoxide in 1902, followed by the discovery of the synthesis of ammonia (NH,) in 1910. The Fischer-Tropsch synthesis was developed in the following decade and then came the manufacture of methanol and higher alcohols, Syngas is now used in a host of different ways. The uses and reactions of syngas in the light of changes that have taken place from 1984 to 1994, with an update, constitute the body of this report. The chemistry of the separate components of syngas, H, and CO, will be discussed in some detail. It is remarkable that, although these are apparently simple diatomic molecules, they are readily adsorbed in diverse ways on the surface of catalysts. Together with carbon dioxide, CO,, they form an unexpectedly large number of complex species with both heterogeneous and homogeneous catalysts. The principal fuel uses of syngas in 1994 are given in Fig. 1. The major commercial, near commercial and potentially commercial chemical uses of syngas in 1994 are outlined in Fig. 2. Wender and Seshadri (1984) summarized the fuel and chemical uses of syngas in 1984 (Figs. 3 and 4). Although it appears from a comparison of the Gasoline Diesel Chemicals

FISCHER-TROPSCH -

METHANOL -

isobulylene

@ME)

Medium BTU GAS (Turbine fuel. ICCC)

Gasoline

Fig. 1. Principal commercial fuel uses of synthesis gas (1994).

MTBE

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189-297

DIMETHYL CARBONATE

FORMIC ACID

ACETIC ANHYDRIDE

ETHYLENE

\ \ CHEMICALS

METHYL ACETATE

WAXES

T

FORMALDEHYDE

\

CHjOll

FC

+ ACETIC ACID

AMMONIAC

CHLORCO I IMETHYL

co .I’ ,’

Rh b’

METHANE5 i

02

AMINES

I

,.’

I

”

Acelaldeh de Ethsno r

Rh

Ethylene

ZSM-5

(hydroformylation)

(Zeolites)

I Co

I

I

J

i

VINYL

OldillS

ALDEHYDES ALCOHOLS

4

ACETATE

Aromatics

(Z-ethylhexanol) ___

Commercial

.

Fig. 2. Commercial,

-.

Near commercial,

near commercial

perhaps availnble -

for license

and potential chemicals

.. .. .. .. . .. ..

Polenlial

from synthesis gas

(next decade)

(I 994).

corresponding figures that there have been no significant new uses of syngas in the past uses of syngas as new decade, there have been great increases in the “traditional” outlets have appeared.

GASOLINE ~_ZBO!l!?r.F,SCHER_TROPSCH~ DIESEL OIL

/ ,*’

1 C,-C6

GASOLINE DIESEL OIL OTHER PRODUCTS

ALCOHOLS

METHANE

MTEIE

H:,

ISOSYNTHESIS

GASOLINE DIESEL OIL AROMATICS

FUEL CELLS

FUEL (neat) (BTX)

____

commercial processes -.-.-.processes commercial III 1984-85 -~~-~~~~~~-~-~~ processes possibly for llle nsxl decade

Fig. 3. Commercial or near commercial (Wender and Seshadri, 1984).

processes

for the production

of liquid fuels from synthesis

gas

I. Weruler/

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Fuel Processing

FORMIC

Technology

ACID

t METHYL FORMATE

AMMONIA

\H2

ETHYLENE GLYCOL

\

i ETHANOL

IXYLENES

Fig. 4. Commercial, 1984).

ACETIC ANHYDRIDE

1

ETHYLENE GLYCOL

-.-.-.-._-......_.....-.

48 (1996) 189-297

cD~_/~~;~A~~~E

\ SINGLE-CELL PROTEIN

1

commercial near commcrciol polenlinl (ncnl decade)

near commercial

and potential

chemicals

from synthesis

gas (Wender

and Seshadri,

Although hydrogen as such is not found in the list of top chemicals produced in the United States, the principal use of syngas is for the manufacture of hydrogen, huge amounts of which are consumed in the synthesis of ammonia. The demand for hydrogen continues to grow. Environmental concerns constitute a driving force for greater growth in the use of hydrogen to produce clean fuels and chemicals. In addition, there has been a spectacular increase in the manufacture of methanol for the synthesis of methyl r-butyl ether (MTBE), which is used in reformulated automobile fuels. Indeed, methanol used for fuel and chemicals has achieved the status of a commodity chemical. The Fischer-Tropsch synthesis remains the third largest consumer of syngas. Sasol, using iron catalysts, continues to expand in the synthesis of fuels and chemicals from coal based syngas. In 1993, the Shell Middle Distillate Synthesis came onstream, utilizing natural gas to produce mainly gasoline and diesel fuel in fixed tubular reactors. Plans for further growth in these newer syngas conversions are contemplated. There is great interest in the utilization of remote natural gas reserves. Exxon has built and operated a large Fischer-Tropsch (FI’) demonstration plant using a multiphase slurry reactor. Both the Shell and the Exxon processes are based on the use of cobalt catalysts. Both produce high molecular weight paraffins (waxes) which are hydroisomerized and cracked to liquid products suitable for refinery or chemical plant feedstocks. Parallel with the development of the Exxon multiphase process, Sasol has developed the so-called Sasol SPD (slurry-phase distillate) process. Until now it is the only commercial experience that has demonstrated successful IT slurry operation. It is directed towards the production of paraffinic waxes; however, the product also constitutes a valuable feedstock for high-quality diesel fuel.

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The hydroformylation of olefins to aldehydes and alcohols, which is completely a chemical use of syngas, remains the fourth largest use of carbon monoxide and hydrogen mixtures (it was inadvertently omitted from Fig. 4). In this versatile reaction, termed the hydroformylation or 0x0 synthesis, a large variety of olefins react with syngas, in the presence of cobalt or rhodium catalysts, to form aldehydes and alcohols with one carbon more than the starting olefin. The wide applicability of this synthesis has resulted in its steady growth in the synthesis of plasticizers, pharmaceuticals and hundreds of other chemicals. As this is a chemical use, with chemicals sold by the pound, the amount of syngas is considerably below that utilized in the production of hydrogen, methanol or Fischer-Tropsch products. A direct use of syngas as a fuel that has potential for rapid growth is as a clean fuel for integrated gasification combined cycle (IGCC) units for the generation of electricity. The amount of syngas employed in this manner may exceed that used in Fischer-Tropsch reactions in a few years. Methanol may play a part as a peaking fuel in IGCC. If excess syngas is available from IGCC units, its conversion to methanol may be the best road to take, for methanol is a convenient storage fuel for peak use. In addition, methanol synthesis plants can be converted to the production of ammonia. A 260 ton per day liquid-phase methanol facility is currently being built in Kingsport, Tennessee. It is the result of a joint venture of the Eastman Chemical Company and Air Products and Chemicals. The process is based on the use of a multiphase slurry reactor for the production of methanol from synthesis gas. Copper based catalysts are being used. The process was developed by Air Products with the sponsorship of the US Department of Energy. It is important to remember that the preparation and clean-up of syngas from any source approaches 60-70% of the cost of the final product. There is continuing steady progress by industrial firms in lowering costs involved in syngas preparation. Mobil’s methanol-to-gasoline (MTG) process is in commercial operation in New Zealand, producing both methanol and high octane gasoline. The plant has been sold to a private company which is a large producer of methanol. The plant has met the strategic goal of reducing New Zealand’s dependence on foreign oil, but methanol could be sold as such rather than being converted to high octane fuel. Mobil has viable processes for the conversion of methanol to gasoline in a slurry reactor and also processes for the conversion of methanol to diesel fuel. Mobil’s ZSM-5 technology has been optimized and the future of these processes would brighten should oil prices rise. A comparison of the data in Fig. 1 with those in Fig. 3 reveals that the commercial syngas processes appear to be the same (solid arrows) but that essentially all of the potential syngas uses (broken arrows) have not come to fruition, except for the conversion of methanol to gasoline via the Mobil MTG process. Although operated in Italy, no plant now exists for the synthesis of higher alcohols directly from syngas; methyl r-butyl and other ethers have taken over. The Isosynthesis deserves further research but conditions are too severe and selectivity is low. Neat methanol use as an automobile fuel is not promising; it picks up water easily and is corrosive to certain automobile parts, among other problems. Ethanol, made by fermentation, is used as an octane extender and as part of ethyl r-butyl ether in reformulated gasoline. The present syntheses of major chemicals from syngas are shown in Fig. 2, which

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should be compared with the diagram shown in Fig. 4. Not too much has changed in terms of outlets for syngas. Ethylene glycol is still synthesized commercially via ethylene oxide; other indirect routes to this glycol are under study. The direct conversion of syngas to ethanol has been studied by many; a two-step synthesis of ethanol via methanol and syngas (the homologation of methanol) is more promising and may become commercial in the next decade. The conversion of methanol to olefins and benzene, toluene and zylene (BTX) via Mobil technology may become of industrial interest in a decade. As a chemical, ammonia still consumes the largest amount of syngas via its conversion to hydrogen. All new plants for the synthesis of acetic acid (derived solely from syngas) will use the rhodium-catalyzed methanol plus carbon monoxide process, which shows steady growth in terms of syngas use. The homogeneous carbonylations of methanol to acetic acid and of methyl acetate to acetic anhydride are, at present, the only examples of syngas use which have displaced petroleum based routes to commodity chemicals. The Eastman Chemical coal based plant for converting methanol to acetic anhydride via acetic acid is a great success and has undergone recent expansion. There is much promise for growth in the production of other oxygenated chemicals from syngas as this plant grows and diversifies. Formaldehyde, long the chief chemical derived from methanol, has been overtaken by the use of methanol for the synthesis of methyl ethers such as MTBE. Mechanistic aspects of most syngas reactions are discussed, but it is not the purpose of the present report to go deeply into this aspect. It is recognized that there remains considerable need for further research in almost every area for the elucidation of reaction intermediates and pathways.

SYNTHE8IS

Hydrogen Methanol Ammonia Carbon Monoxide Medium BTU Gas Methane Higher (C,-Cd Alcohols Gasoline Diesel Fuel Isobutanol

INDLRJXTSYNTHESIS (VIA METHANOLS

Formaldehyde Acetic Acid Methyl Acetate Acetic Anhydride Vinyl Acetate Methyl Formate Formic Acid Ethanol Dimethyl Carbonate Dimethyl Oxalate Gasoline Diesel Fuel Ethylene Propylene m-x Chloromethanes Methylamines Methyl Glycolate Ethylene Glycol Fig. 5. Fuels and chemicals

Olefins H&O Aldehydes Co/Rh - Alcohols Isobutyle-ne Q&Q& H+ Acetylene n Ni Olefms &_ co

MTBE

Aaylic Acid

Highly-branched Acids -

RCOOH +2Q RlQ&

RCH.#OOH

Nimxuomatics a

Isocyanates Pd

Teqhtbalic

Acid CH,OH

from synthesis gas.

Dimethyl Taephthalate

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This report classifies syngas reactions into six categories which, although they may occasionally overlap, are relevant enough to stand apart. The choice of categories rests to a great extent on the amount of syngas used and the commercial importance of the products. The categories are listed below. 1. Synthesis gas as a source of hydrogen. 2. Direct conversion of syngas to fuels and chemicals. 3. The hydroformylation (0x0) reaction. 4. The Mobil methanol-to-gasoline (MTG) and related processes. 5. Methanol plus syngas or carbon monoxide for the synthesis of chemicals. 6. Miscellaneous reactions. A summary of fuels and chemicals obtained from syngas reactions is given in Fig. 5.

1. Introduction I .I. Background The main body of this review will not deal in detail with the manufacture of syngas and of the separation of hydrogen and of carbon monoxide from syngas. Various sources and processes for syngas production yield different ratios of H, to CO and the availability of source material for syngas manufacture varies at different places in the world. Matching the specific method of production of syngas with the use in a particular synthesis is an important factor in determining the technical and economic value of the various processes (Haag et al., 1987). Natural gas is the largest source of syngas at present and its use for this purpose is growing. Methane is the chief constituent by far of natural gas. Petroleum fractions are the next largest syngas source at present and significant quantities of syngas are being made from coal. Thermodynamically, methane is an exceptionally stable molecule, but it can be activated by reaction with such reagents as oxygen or water. A great deal of research has been carried out on the reaction of methane with oxygen but total oxidation to CO, is usually the major product. The equilibrium constant for the reaction of methane with water, the well-known steam reforming reaction, is more favorable, with an equilibrium constant of unity at 627°C. The conversion of methane to syngas is carried out commercially on a very large scale in the presence of a nickel catalyst at about 700°C under pressure. The CO, concentration is reduced by the presence of H, in the reaction product. As Notari (1991) has pointed out, “ALL industrial production for which methane is the raw material proceeds through the intermediate production of syngas: it is the ONLY reaction by which methane can be transformed into reaction intermediates (H, and CO) with limited amounts of undesired CO,.” The route to methanol, which is made commercially in large plants throughout the world, is through syngas in essentially all cases, with natural gas as the chief source of the syngas. Since natural gas has been in short supply in Europe, petroleum fractions have been used as a source of syngas. The near-panic programs pursued following the

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Fuel Processing Technology 48 (1996) 189-297

1973-1974 oil embargo involved coal as the source of fuels and chemicals, including increased gasification of coal to syngas. Coal utilization in the United States and worldwide increased and this increase has been sustained, but coal research and development have decreased with the decline in oil prices and the increased availability of natural gas. The long-term instability of oil and gas markets is a reason for sustaining research, development and technology for the use of syngas (often referred to as C-l chemistry). There are a few salient points, mostly well known, that apply to the supply of fuels and chemicals in general and to syngas in particular: (i) perhaps the most precarious situation involves the availability of petroleum, as the largest oil resources exist in politically unstable areas far from the United States. In terms of availability, natural gas and coal constitute large resources and are ubiquitous. (ii> Petroleum, natural gas and coal will continue to compete for world markets for several decades. Flexibility will be a key so that a particular application will not depend on the availability of one kind of fuel. As all of these resources are convertible to syngas, applications based on syngas use will continue to grow. (iii) It is not possible to separate the fuels and chemicals businesses as some are prone to do. The use of fuels will govern the use of chemicals, as is the case in the petroleum industry. Fuels are sold by the barrel or by the ton; chemicals are of higher intrinsic value and are sold by the pound. (iv) With regard to coal use, gasification processes are more forgiving than direct liquefaction. Variability in coal source, rank, mineral matter, etc. are more easily accommodated by gasification processes. Indirect liquefaction (via syngas) products are relatively easily converted to clean gaseous and liquid fuels and chemicals. (v) The major cost of almost all uses of syngas lies in the gasification and purification of the syngas. (vi) Syngas will continue to be the source of methanol, although there is significant ongoing research on the direct oxidation of methane to methanol and products derived from methanol, such as formaldehyde. (vii) Syngas will continue to be the world’s principal source of hydrogen. Hydrogen has a great many uses, including primary application for the manufacture of ammonia (NH,). (viii) Environmental and economic factors are the eventual determinants in fossil fuel use. Even a large amount of an available resource, if environmentally harmful, will tend to lose out sooner or later. Environmental constraints are found in gas clean-up and in the nature of gasification residues. The latter should be characterized, be non-leachable and, if possible, be useful in some way. (ix) It is almost inevitable that future power plants will allow the use of any clean fluid fuel: natural gas, oil, syngas, petroleum, methanol, or even a Fischer-Tropsch derived liquid. There is a growing tendency for coal and petroleum coke to be converted to syngas and then used in a type of integrated gasification combined cycle (IGCC) mode, as first demonstrated at Cool Water, California. This procedure could result in the proliferation (often nearby) of plants that convert the syngas to high-quality transportation fuels or to chemicals, or use the gas as a clean industrial fuel. (XI The modem petrochemical industry is based on natural gas and the by-products of petroleum refining. The so-called petrochemical feedstocks are comprised of ethylene, propylene, the butenes, benzene, toluene, and the xylenes. All of these and liquid fuels can be made from syngas, and it is therefore possible to envision a petroleum-less refinery. (xi) As tetraethyllead disappears from automobile gasoline tanks, the demand will grow for oxygenated chemicals with high

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octane ratings. Economical syntheses of such fuel-blending agents as MTBE, other ethers or higher alcohols are needed. (xii) Sulfur and nitrogen compounds are removed from crude syngas and can be sold as solid sulfur or as ammonia, respectively. (xiii) A major factor is the discovery of large natural gas supplies. The use of natural gas as a source of syngas has the advantage of producing less carbon dioxide, a greenhouse gas, than syngas from other sources. It is considerably cheaper to obtain syngas from natural gas than from naphtha or coal. (xiv) Fossil fuels will remain the dominant source for the production of energy; technologies that are economically feasible and minimize pollution must be developed. There are many issues associated with the integration of gasifiers with syngas conversion processes (Haag et al., 1987). The downstream conversion process dictates a minimum partial pressure of the active compounds (HZ and CO) in the syngas. This minimum partial pressure is affected by three factors in the gasifier: gasifier pressure, purity of the oxygen used, and the amount of light paraffins (mainly methane) generated in the gasifiers. In addition, the syngas pressure is changed by using either a gas compressor or an expander. The HZ/CO ratio of the syngas is another important issue. It is costly to decrease the HZ/CO ratio because of thermodynamic constraints. However, the use of a water gas shift (WGS) unit to increase the HJCO ratio is a proven technology that adds substantially to the process cost. Furthermore, syngas conversion processes use catalysts that usually require syngas with less than a defined minimum threshold level of impurities (usually H,S and COS). Expensive syngas purification units are needed to remove excess impurities. Another important integration issue is the co-production of energies in various forms. There are actually two sub-issues involved here. One is the co-production of fuel gas and syngas with liquid fuels to achieve high thermal efficiency. In a coal based synfuel plant, it is almost always more costly to have a process designed to produce only liquid fuels. By recycling the light hydrocarbons to make additional syngas, 3040% of the energy in the hydrocarbons is easily lost. The other sub-issue is the co-production of steam and electricity, as in cogeneration. With proper integration and optimization of a gasifier with a syngas conversion process, the resulting high thermal efficiency can result in a surplus of steam or electricity. The steam may be exported to nearby factories as process steam or to nearby communities as utility steam, and the electricity is exported to the local utility network. 1.2. H, + CO partial pressure requirements The degree of importance attached to high H, + CO partial pressures in the syngas depends strongly on the syngas conversion process. Molar contraction almost always occurs with syngas conversion; consequently, conversions are thermodynamically more favorable at higher H, + CO partial pressures. However, the importance of this requirement varies with the conversion process. Generally speaking, conversion reactions that are limited by chemical equilibrium, such as the methanol synthesis, will require higher partial pressures. Fischer-Tropsch (IT) reactions are less strongly constrained by equilibrium considerations and can tolerate low syngas partial pressures.

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In commercial application, the synthesis of methanol requires a minimum syngas pressure of 600-700 psig. Although FT reactions may proceed to high syngas conversions at low pressures, moderate to high partial pressures (20-1000 psig) are usually used to improve process economics. The easiest way to obtain the required syngas partial pressure is to select gasifiers that operate at the proper pressure. If a low-pressure gasifier is employed, the syngas is compressed to the required value. It is more efficient to compress gasifier oxygen and use high-pressure steam for gasification than to compress output syngas. The penalty for gas compression is relatively smaller at high pressures (e.g., above 300 psig) because the compression ratio is reduced. The best operating gasifier pressure ranges from = 300 to 600 psig. Most impurities in syngas gasifiers are removed fairly easily. Removal of sulfur compounds is essential because they poison the catalysts used in the downstream processes. Established purification processes for removal of sulfur compounds have been given by Kuo (1984). The maximum allowable sulfur levels for the synthesis of methanol and for the Fischer-Tropsch reaction are about l-2 ppm by volume. As mentioned earlier, different ratios of Hz/CO are obtained from various feedstocks. From natural gas CH,+H,Oe3H,+COAH=

+49.0kcalmol-’

(1.1)

From oil fractions -CH,

- +H,O ZZ2H, + CO AH= +36.1 kcalmol-’

(1.2)

From coal C+H,Ot3H,+COAH=

+3l.lkcalmol-’

( 1.3)

The energy required for these processes can be supplied by external heating or by

partial combustion of the feedstock. At present, a great deal of research and development is aimed at reducing the cost of producing and purifying syngas, as these operations constitute the major cost in the conversion of syngas to fuels and chemicals (Rostrop-Nielsen, 1994, Sunset et al., 1994). Syngas and its separate components are involved in a great number of reactions and chemical syntheses. It will be of help to memory, and to devising new chemical syntheses, to separate syngas conversions into presently commercial, potentially commercial, and processes which have been investigated but, at present, do not appear to be able to replace existing processes not based on syngas. Syngas conversion chemistry has undergone improvement over the last decade but, as will be shown, many attempts to convert syngas directly to important uses, especially to chemicals, have not been successful as yet. 1.3. Carbon monoxide: structure and reactivity Before discussing syngas reactions, a short consideration of the structure and reactivity of the carbon monoxide molecule is in order. The CO molecule contains a

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total of 14 electrons, six contributed by carbon and eight by oxygen. Its structure may be represented as :C:::O:, in which only the valence electrons are depicted. Oxygen transfers an electron to carbon producing a structure with three bonds; extra stability is provided by coulombic interaction of the oppositely charged atoms. One of the principal arguments in favor of a triple bond in carbon monoxide is its analogy with the nitrogen molecule which also has 14 electrons and unequivocally can be assigned a triple bond structure. Indeed there is a striking similarity between the physical properties of CO and N, (Orchin and Wender, 1957). The bond dissociation energy of N, is 225.1 kcal, and that of CO is 255.7 kcal. This is consistent with the triple bond formulation of CO. The bond energy of CO is the largest observed bond energy of any diatomic molecule; the CO molecule does not dissociate in the synthesis of methanol or in many carbonylation reactions. It is therefore remarkable that CO does dissociate readily over certain transition metals. This is a necessary condition that must take place on a Fischer-Tropsch catalyst. There are important variables which are necessary for this to take place; occupation of the CO 2n (antibonding) orbital of CO and bonding of the dissociated carbon and oxygen atoms to transition metal catalysts. Dissociation of CO also occurs in the methanation of syngas, in the Boudouard reaction and in the conversion of syngas to higher alcohols over heterogenous catalysts. Bonding in metal-CO complexes and adsorption on transition metals involve both o and rr components. Transition elements readily form strong bonds with compounds (CO, ethylene) that contain n-electron systems or which have orbitals of suitable energy and symmetry to form dn-bonds. Bonding between transition metal atoms and CO is of great importance because transition metals, deposited on supports or as discrete complexes, are catalysts for the reaction between CO and most other molecules. Multiple bond character in the metal-carbon bond occurs by formation of metal-CO rr bonds by overlap of metal dr orbitals with empty antibonding CO orbitals (Fig. 6). Hydrogen plays at least two important roles in syngas reactions (Bartholomew, 1988). It acts as a reductant to activate the catalyst; this affects the initial chemical catalyst states in ensuing reactions. Secondly, as H, is a reactant in the hydrogenation of CO, its adsorption kinetics, energetics and interactions with CO/carbon surface species greatly affect the activity and selectivity behavior of the catalysts used. Carbon monoxide reacts as a weak base (an electron donor) in many acid-catalyzed

filled

metal

d orbital

V empty v*

CO

orbital

Fig. 6. Molecular orbital representation of a carbon monoxide transition metal bond.

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reactions, e.g., with carbenium ions: R+ +:C=O --) (RCO)+. It reacts less readily as an electron acceptor, as in the synthesis of formates; here CO acts as an electron acceptor, with an empty orbital on the carbon atom: B- + C=O + (BCOl-. There are four commercial processes for the purification of CO. Two are based on CO absorption by salt solutions, one uses low temperature condensation or fractionation and the fourth involves adsorption of CO on a solid adsorbent material (Pierantozzi, 1993). Pressure swing adsorption on high-area materials is being adopted increasingly in preference to cryogenic separation. Similar techniques to remove impurities are used in all four processes.

2. Classification

of syngas reactions

2.1. Major commercial uses of syngas This report will classify syngas reactions into categories which, although they may occasionally overlap or seem to intrude on one another, are relevant enough to stand apart. The choice of the categories rests to a great extent on the amount of syngas used and the commercial importance of the products. Another factor is whether the syngas is converted directly to fuels and chemicals (direct synthesis to products) or whether the product of the direct synthesis reacts further with either syngas or carbon monoxide (i.e. syngas is first converted to methanol which then reacts with carbon monoxide to form acetic acid). The largest commercial use of syngas is in the manufacture of hydrogen gas, more than half of which is used in the synthesis of ammonia. The second largest commercial use of syngas is in the synthesis of methanol. The third largest industrial use of syngas is in its conversion to paraffins, olefins and oxygenates via the Fischer-Tropsch reaction. There is recent renewed interest in this synthesis with the start of operation of new plants. These three uses account for the main industrial commercial utilization of syngas. The fourth largest industrial use of syngas, the hydroformylation (0x01 reaction, although considerably smaller than the uses given above, is a viable, flexible reaction that is used in a large number of syntheses of important industrial products. Hydroformylation consists of a general and widely applied conversion of olefins plus syngas to aldehydes and alcohols. It is often used to synthesize intermediates in a sequence of synthetic reactions. Hydroformylation plants are found in many countries throughout the world. The use of syngas in the generation of electricity via integrated gasification combined cycle (IGCC) from coal and petroleum coke has the potential for considerable growth. 2.2. Categories of syngas reactions Category 1: syngas as a source of hydrogen. Category 2: direct conversion of syngas to fuels and chemicals: of methanol and the Fischer-Tropsch reaction.

includes the synthesis

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Category 3: the hydroformylation (0x0) reaction: olefins plus syngas to aldehydes and alcohols. Category 4: the Mobil methanol-to-gasoline (MTG) and related processes. Category 5: methanol plus syngas or carbon monoxide, for the synthesis of chemicals. Category 6: miscellaneous carbonylation reactions.

3. Category

1. Manufacture

of hydrogen

Syngas is our chief source of hydrogen gas. When syngas is treated with water in the presence of generally nickel based catalysts, the CO reacts via the water-gas shift (WGS) reaction to yield a second mole of hydrogen, if required. The WGS reaction is also used to obtain hydrogen to carbon monoxide ratios desired for particular syntheses: about 2H,: ICO for the synthesis of methanol and for the Fischer-Tropsch reaction and about 3H,: 1CO for the synthesis of methane. Approximately ten trillion (1012> standard cubic feet (SCF) of hydrogen are consumed annually in the world; about four are used per year in the United States. Worldwide, most hydrogen is used for the production of ammonia (almost six trillion SCF of syngas per year). Refinery needs for H, have been expanding rapidly, meaning significant growth in hydrogen demand. New environmental regulations are driving more demand for hydrogen for better gasoline quality and improved diesel. Another very large growth in demand for hydrogen will be for process needs and syngas (for methanol) to satisfy the enormous growth in MTBE use. Detailed uses of hydrogen as a basic material for large-scale processes are: ammonia synthesis, manufacture of methanol, hydroformylation of olefins to aldehydes and alcohols (the 0x0 reaction), hydrogenation to form organic intermediates (amines, cyclohexane, aliphatic alcohols and natural fats and oils), iron ore reduction, protection gas, petroleum processing (hydrocracking, hydrodealkylation, hydrodenitrogenation, hydrodesulphurization, and hydrorefining). Fischer-Tropsch synthesis, methanation, coal hydrogenation, heavy oil hydrogenation, in rocket propulsion, etc. Hydrogen is also expected to play a major role as an energy carrier. Estimates of future hydrogen demand in the non-energetic area are outlined by the assumed growth rates of various products. Specifically, hydrogen demand for synthesis processes is likely to grow because of the already existing high level of technological development. Environmental regulations are in part responsible for increased hydrogen demand. Hydrogen used in the chemical and petrochemical industries is produced primarily from fossil energy carriers (natural gas, crude oil, and coal) and steam. Water electrolysis to produce hydrogen is restricted to areas with low hydrogen demand, since electric power costs are almost prohibitive. However, significant amounts of hydrogen from water electrolysis for ammonia synthesis are produced in a few countries with sufficient electric power from hydroelectric plants. The principal commercial processes for hydrogen manufacture from fossil energy carriers are catalytic steam reforming, pressure partial oxidation of hydrocarbons, and coal gasification. In the United States, most industrial hydrogen is manufactured by steam reforming of natural gas. Relatively small

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amounts of hydrogen are produced by steam reforming of naphtha and by partial oxidation of oil. Processing difficulties and manufacturing costs increase as the feed is changed from natural gas to liquid hydrocarbons and then to coal; partial oxidation and coal gasification processes require higher capital costs. Because of the newly created hydrogen demand for the gasoline and diesel pool, improved reformer designs are underway to reduce the cost of hydrogen production. A cheap way of producing hydrogen from water is a most desirable goal. The product gas leaving the hydrogen manufacture unit from fossil energy carriers is a mixture of hydrogen, carbon monoxide, carbon dioxide and small amounts of methane. Purification steps of hydrogen include WGS catalysis, CO, removal, methanation, pressure swing adsorption, and cryogenic separation. The CO is converted to CO, and hydrogen via the WGS reaction (CO + H,O + CO, + H,) which has the advantage of increasing the hydrogen yield at the expense of carbon monoxide. Carbon dioxide is scrubbed out by ethanolamine solutions or other processes such as Rectisol, Selexol, Sulfinol, amine guard or hot potassium carbonate scrubbing. Remaining small amounts of carbon oxides are converted to methane over supported nickel methanation catalysts. Pressure swing adsorption (PSA) is a single step process in which all the constituents, except hydrogen, are simultaneously removed as the reformer effluent moves through the adsorbent bed. In a cryogenic purification system, the feed is cooled by indirect heat exchange and impurities are condensed and removed as liquid by-products. The chief commercial uses of syngas are the following in decreasing order: as a source of molecular hydrogen > in the synthesis of methanol > Fischer-Tropsch synthesis > in the hydroformylation of olefins to aldehydes and alcohols > in carbonylation reactions in which carbon monoxide reacts with other molecular species. There is a large potential for a rapidly growing use of syngas as a fuel for electricity generation via integrated gasification combined cycle (IGCC).

4. Category 2. Direct conversion of syngas to fuels and chemicals 4.1. Background and scope There are five major direct uses of syngas which do not involve a third chemical reactant, i.e. syngas is used as such. They are: - Synthesis of methanol-commercial - Fischer-Tropsch and related reactions-commercial . The Isosynthesis-not commercial * Direct combustion for generation of heat and/or electricity-commercial * Methanation of syngas to synthetic natural gas (SNG)-one plant in the US. The synthesis of methanol and the Fischer-Tropsch synthesis (FTS) are the two major direct uses of syngas for the production of fuels and chemicals. Each will be discussed separately, but it is of considerable interest to point out similarities and differences of these two important syntheses. The original patents for both syntheses were obtained by BASF in Germany in the same year, 1913, (Anderson, 1984; Lee, 1990). The first commercial plant for the conversion of syngas to methanol was built in

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1923 by BASF. The first commercial plant for the FTS began operation in Germany in 1935. Both syntheses are conducted essentially with a syngas ratio of 2H,: ICO. The methanol synthesis is equilibrium limited; the ITS is limited by AndersonSchulz-Flory selectivity and kinetics. The methanol synthesis, as practiced commercially, yields methanol in over 99% selectivity; the FIS yields a multitude of products (and incidentally, very little methanol) Transition metal catalysts are used in both syntheses. Methanol synthesis catalysts, found in the right side of the periodic table, do not split the carbon-oxygen bond in carbon monoxide. In a basic difference, FT catalysts, found in the left side of the periodic table, proceed mainly by actual breaking (dissociation) of the carbon-oxygen bond in carbon monoxide. The weight retention of syngas (2H, + lCO> as a feedstock for methanol approaches 100%. In the FT reaction with iron catalysts, each methylene (-CH,-) group in the formed carbon chain is accompanied by formation of water or carbon dioxide depending on the catalyst promoter level and operating temperature. With a cobalt catalyst in the FT process, water rather than carbon dioxide is formed. Methanol is the source material for a large number of chemicals, mostly oxygenated compounds. Many of the chemicals involve reaction of methanol with one or more other compounds. A large variety of chemicals is also produced in the FI’S but complex and sophisticated separations and upgrading schemes are necessary to obtain these chemicals. Methanol, when first produced from syngas, was used almost entirely as a chemical and chemical precursor. In the last decade and in the upcoming years, the fuel uses of methanol, largely to form methyl ethers, will overtake its use for chemical production. The FT synthesis began as a process to produce liquid fuels. In the last decade, however, the production of chemicals and of widely used waxes via the FT reaction has grown in importance. Methanol use has expanded rapidly as a component (35.3 percent) of methyl tert-butyl ether (MTBE), and other ethers which are oxygenated octane enhancers and fuels. The FIS has new potential in the synthesis of high molecular weight products which are then cracked to yield high-grade diesel fuel in addition to gasoline. 4.2. Methanol manufacture 4.2. I. Introduction The world consumption of methanol is over 24 million tons per year and is still

growing (Gee& et al., 1990). Methanol is one of the top ten organic chemicals manufactured in the world. The alcohol is a fuel itself, an octane extender, it is used in the manufacture of other fuels such as methyl t-butyl ether and higher alcohols, and it is a precursor to a large and varied number of high-value chemicals and polymers. The

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many uses of methanol will be discussed later. Almost all methanol is synthesized from syngas containing 2-8 ~01% of CO, (H,/CO/CO,) derived to a very large extent by steam reforming of natural gas. Small amounts of syngas are obtained by steam reforming of naphtha and even smaller amounts by gasification of coal. The ratios of H&O/CO, depend on the source. Methanol is catalytically synthesized by the reactions CO+2H,+CH,OH,

AH&,=

AG&. CO, + 3H, --) CH,OH + H,O,

-24.0kcalmoll’ = + 10.8 kcal mol- ’

(4.1)

AH;,,., = - 14.7 kcalmol-’ AG&,oC = + 14.8kcalmoll’

(4.2)

The water gas shift (WGS) reaction occurs simultaneously with methanol synthesis. CO + H,O + CO, + H,

AH~,,.,

= -9.3

kcalmol-’

AG”,,,., = -3.9kcalmol-’

(4.3)

The synthesis of methanol from syngas is one of the technically very well-developed industrial processes. Although a great deal of information has been published on its chemistry, the mechanism of the commercial synthesis of methanol from syngas has been the subject of much controversy. In France, Patart (1921) obtained a patent covering the hydrogenation of CO to methanol and other oxygenated compounds at elevated temperatures and pressures with catalysts containing Cr, Co and Mn in the metallic form in addition to various oxides and other compounds. In 1923, BASF obtained a patent for the synthesis of methanol and built the first commercial plant for the conversion of syngas to methanol using a zinc oxide/chromium catalyst (BASF, 1923; Lormand, 1925). As recently as 1923, methanol was produced by the distillation of wood. How many trees have been saved by the development of catalysts for the methanol synthesis-ecological benefits through catalysis (Waugh, 1992)! At this stage, realizing that methanol could be manufactured more economically catalytically, DuPont and the Commercial Solvents Corporation started experimenting with the catalytic synthesis of methanol and, in 1927, commercial production by the high-pressure process began in the United States (Wade et al., 1981). In the Commercial Solvents process, methanol was produced from a HZ/CO, feedstock obtained from a butanol fermentation plant at 30.6 MPa with metal oxide catalysts (Lee, 1990). The BASF process used a high temperature catalyst operating at 3.50-400°C because of its low catalytic activity. These catalysts had to operate at high pressures (25-35 MPa) because of low syngas conversion resulting from less favorable thermodynamic equilibrium limitations at high temperatures. The catalysts are less active than copperzinc based catalysts but are more tolerant towards poisoning by compounds containing sulfur.

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ICI introduced the more active copper-zinc based catalysts in 1966 (Davies et al., 1966). Researchers in the 1920s had found that copper was the key component for the methanol synthesis; newer developments made possible the use of copper/zinc oxide catalysts. The source of syngas had shifted from coal to liquid or gaseous hydrocarbons such as steam reforming of methane or naphtha. The syngas feed could be made pure enough for copper catalysts (sulfur and arsenic compounds were now removed in addition to chlorine). The copper-zinc oxide catalyst can be permanently deactivated at high temperatures, so that proper control of reactor temperature is necessary. However, even with the most carefully prepared catalysts, small amounts of methane, dimethyl ether and traces of higher alcohols appear among the products. ICI now had a new generation of low-pressure (< 10 MPa) and low-temperature (220-270°C) Cu/ZnO/Al,O,-catalyzed methanol synthesis established as a commercial process. A methanol process based on a Cu/ZnO catalyst was put into operation by Lurgi in 1971; Haldor Topsije and BASF also practice low-pressure methanol synthesis using the Cu/ZnO/M,O, (M is Al or Cr) catalysts. A basic difference between the ICI and the Lurgi process is the type of reactor used; ICI uses multi-quench reactors and Lurgi employs multi-tubular reactors. Modem methanol plants can yield about 1 kg of methanol per liter of catalyst per hour, stated to be the “magic figure” for the low-pressure process (Herman, 199 1). The modem synthesis of methanol is very selective; > 99.5% conversion to methanol is obtained. This high selectivity is quite an achievement in view of the fact that methanol is thermodynamically the least probable product of the syngas conversion, i.e., other compounds are formed with a more negative free energy change than methanol. A graphical example is given in Fig. 7, which shows the standard Gibbs free energy change at 600°K (327°C) in kcal mol- ’ of carbon as a function of chain length n in the product alcohols or paraffins (Wender and Klier, 1989). nC0 + 2nH,

+ CH,(CH,),_

(2n - 1)CO + (n + l)H,

,OH + (n - l)H,O --j CH,(CH,)np

(4.4)

,OH + (n - l)CO,

(4.5)

nC0 + (2n + l)H,

-+ CH,(CH,),,_,CH3

+ n&O

( 4.6)

2nCO+(n+

-+ CH,(CH,),,_,CH,

+ nCO,

(4.7)

l)H,

For long chains, alcohols plus water formed in Eq. (4.4) tend to have the same standard free energy as the paraffins plus water of Eq. (4.6); alcohols plus CO, formed in Eq. (4.5) have the same standard free energy as paraffins plus CO, in Eq. (4.7). The difference for any single reaction type with water or CO, coproduct is given by the standard free energy of the WGS. For short chains, the free energies of formation for alcohols and hydrocarbons from CO/H, diverge, with hydrocarbons being significantly more favored. The greatest thermodynamic driving force is for production of methane plus CO, and the least thermodynamic driving force, given by a positive Gibbs free energy, is towards methanol. Because of the negative volume change of Eq. (4.1) and Eq. (4.21, the methanol synthesis can be thermodynamically driven against positive free energy by

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GIBBS FREE ENERGY OF CARBON MONOXIDE HYDROGENATION TO

PARAFFINS + CO1

I

;

9

I

I

,

1

,

!

I

3

4

5

6

7

6

5

IO

,

NUMBER OF CARBON ATOMS IN THE HYDROGENATED PRODUCT

Fig. 7. Gibbs free energy of hydrogenation of carbon monoxide.

high pressures, but the catalyst must kinetically prevent formation of all thermodynamically more favored products, i.e., hydrocarbons and C, + alcohols. 4.2.2. Existing technology Methanol is considered a rather cheap chemical (about 50 cents per gallon in July 1996). New plants are being constructed all over the world. Significant new or revamped annual capacity of neat methanol will be completed, engineered, or planned in the 1995-1998 period and beyond, evidently with confidence in the future world market for methanol in fuel uses. Methanol synthesis technology based on steam reforming of natural gas has advanced to the point where room for further gains in energy efficiency is small. Overall efficiency, because of continuing development, has risen from 58% in early plants to 72% with current designs (Abbott et al., 1989). But the added extra heat recovery equipment has been accompanied by rising capital costs. World-scale methanol plants are capital intensive, which has a large influence on production costs. The objective of major producers of methanol is to reduce capital costs significantly without loss of efficiency. A flowsheet for a world-scale methanol plant with the distribution of capital costs allocated to the principal process stages has been outlined by workers at ICI (Fig. 8). Other methanol producers have similar flowsheets. Steps involving desulfurization and distillation will not yield large savings. The focus of development for the past 15 years by methanol producers is in the stage of manufacture of the synthesis gas. Over half the

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Processing

Reforming/ Gas Cooling

Technology 48 (1996) 189-297

Compression

Distillation

Synthesis

32%

207

22%

6% t-ii,

I FFEDSM)CK

PRODUCT

r-27-t

Steam

0

-----

i

Raising

Fig. 8. ICI methanol plant capital cost breakdown

(ICI brochure).

investment in process equipment, with considerable cost-saving potential, is in the reforming stage, with its associated gas cooling and steam raising equipment. We shall deal with these aspects of the methanol synthesis only in passing. A new proposed liquid-phase system for the production of methanol is being built in Kingsport, TN, in the Eastman Chemical Co. facilities. This is the result of a joint venture between Eastman and the developers of the technology, Air Products and Chemicals. This is an example of the commercialization of a project co-sponsored with the US Department of Energy. Use of a slurry bubble column instead of a multi-tubular bed significantly reduces the capital cost of the reactor and decreases the compression cost by reducing the pressure drop across the reactor. The fact that it is a fluidized reactor allows better control of the reaction temperature, avoiding temperature runaways and cold spots. Although this technology has been proven in different situations at the LaPorte, TX facilities by Air Products, continuous operation would be the test that methanol producers may be waiting to see. The ICI and Lurgi low-pressure technologies are briefly discussed below. The earlier high-pressure technologies are omitted in the present review because all new methanol plants built after 1967 utilize the more modem low-pressure technology. The service life of modem ICI catalysts is 3-4 years of continuous operation. Two large ICI methanol plants are a part of the methanol-to-gasoline (MTG) complex in New Zealand. Plants ranging in size from 47000 to 580000 tons per year are in operation, and more are being built. A methanol plant using spherical reactors reported to weigh about one-half as much as the usual cylindrical reaction vessels containing the same catalyst volume is now in operation in Chile with a capacity of 2500 tons per day, possibly the world’s largest single-train methanol synthesis plant (LeBlanc and Rovner, 1990). A methanol synthesis plant producing 750000 tons per year went on stream in Venezuela in 1994 (Hydrocarbon Processing, 1992). In most existing plants, natural gas is steam-reformed to hydrogen-rich syngas, HJ(2CO + 3C0,) > 1. Alternatively, the naphtha reforming process yields a nearly stoichiometric syngas, H,/(2CO + 3C0,) = 1. Coal or heavy oil can be partially oxidized to syngas rich in carbon, HJ(2CO + 3C0,) < 1, which contains considerable quantities of sulfur. Because copper based catalysts are extremely sensitive to sulfur poisoning, the coal-derived syngas must be purified to bring the sulfur content below 0.1 ppm. This can be achieved by several purification processes which operate by physical or chemical adsorption of acid gases, followed by a catalytic purification stage. The

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Crude

Fig. 9. Diagram of a modem plant for methanol production from natural gas. (Reprinted Chinchen et al., copyright 1990, American Chemical Society.)

with permission

from

adjustment of hydrogen-to-carbon ratios suitable for methanol synthesis can be achieved by the WGS reaction between hydrogen-poor syngas and steam. A flow diagram for a typical low-pressure methanol synthesis plant from natural gas is shown in Fig. 9 (Chinchen et al., 1990). In the ICI process, the synthesis loop contains a circulator, a converter, and a heat exchanger. There is a single catalyst bed with lozenge distributors for the injection of cold quench gas located at optimal depths of the catalyst bed. Good mixing of gases and temperature distribution in the reactor are ensured by this design. The distillation plant consists of a unit that removes volatile impurities such as dimethyl ether, esters, ketones, and iron carbonyl, and a unit which removes water and higher alcohols. After the first Billingham methanol plant was operated at 5 MPa since 1966, the pressure of 10 MPa was selected for a second, larger plant, with the carbon efficiency, defined as 100 X (mol of methanol produced)/(mol of CO + CO, in the synthesis gas), 17% percent higher than that of the 5 MPa process. Pinto and Rogerson (1977) have pointed out, however, that the above pressure advantage in efficiency holds only for hydrogen-rich syngas from natural gas or naphtha and not for coal-derived carbon monoxide-rich syngas. Coal gasification conditions are usually such that syngas with high CO/CO, ratios is obtained, which results in high carbon efficiencies over a wide range of pressures in the synthesis loop. Methanol synthesis is adaptable, without loss of carbon efficiency, to match a range of output pressures from various coal gasifiers. The general range of operating conditions of ICI low-pressure methanol plants is 5-10 MPa, 220-280°C

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209

GHSV 5000-60000, and H,/(2CO + 3C0,) ratios 2 1 but adaptable to < 1. Economic considerations have to take into account the energy and capital costs; an advantage associated with the use of coal-derived syngas is that the syngas compression can be avoided without penalty in carbon efficiency. Although Cu/ZnO/Al,O, catalysts have been optimized for maximum selectivity, space time yields, and long service life, they still deactivate, but detailed reports on the rate of deactivation are not available. The main deactivation mechanism probably involves poisoning by chemical impurities, includin g traces of iron carbonyl, even in copper-lined reactors with pre-purified gases. The Lurgi low-pressure technology also utilizes copper based catalysts, principally Cu/ZnO/CrzO,, the detailed preparation and additional promoter composition of which are not disclosed. The Lurgi reactor is a multi-tubular type, the tubes being filled with the catalyst and cooled with pressurized boiling water on the outside. The reactor operates at 240-270°C and good heat transfer to the pressurized boiling water (which yields steam) is achieved. Methanol is produced at space time yields (STYs) close to 1 kg of methanol per liter of catalyst per hour. The crude methanol product is condensed, cooled, and distilled. The reactor has been called “quasi-isothermal” because of the smooth temperature profile along the reactor tubes, with exotherm variations not exceeding 10°C (Supp, 1981). Catalyst life is 3-4 years. The process can be easily adapted to utilize coal-derived syngas by replacing the steam and autoreformers by coal gasifiers and a purification plant, mainly to remove sulfur from the syngas. In the low-pressure industrial plants, the feed composition is adjusted to contain 4-8% CO, in addition to H, and CO. Only 4-7 ~01% methanol is obtained per pass; the remaining gas is recycled. Flow sheets outlining the Nissui-Tops&, the Mitsubishi Gas Chemical and the Japan Gas-Chemical Company Methanol synthesis processes are to be found in the book by Lee (1990). 4.2.3. On the mechanism of the methanol synthesis Before discussing other catalytic systems for the synthesis of methanol, it is worthwhile to examine the mechanism of the commercial system using the Cu/ZnO/Al,O, catalyst. The selective synthesis of methanol over this copper-containing catalyst is a well-developed technology, but there are a number of scientific issues that pervade the literature to this day. Questions include the nature of the active catalytic centers, the reactive carbon-containing component (CO or CO,) and the overall mechanism of the methanol synthesis. The Cu/ZnO/AI,O, catalyst will make methanol in pure H/CO mixtures or when only H,/CO, is used. Most workers in the West assumed that methanol was a product of the reaction of CO in H,/CO/CO, mixtures and developed kinetics and mechanisms based on this premise (Kung, 1980; Klier, 1982; Bart and Sneeden, 1987; Bridger and Spencer, 1989; Chinchen et al., 1988). Russian workers (Kagan et al., 1975; Rozovskii, 1980) however, obtained evidence that methanol was synthesized from CO, and that little or no methanol was formed from H,/CO mixtures with a Cu/ZnO/Al,O, catalyst. A definite answer cannot be obtained from reaction kinetics because the WGS reaction takes place during the synthesis of methanol, so that CO and CO, can

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soecllic Radioaclivity K/m mot 0.14Inlet

CO2

0.12-

1ox;04

5XiO4 Swce

Velocity (h “1

Fig. 10. Specific radioactivities of carbon monoxide, carbon dioxide and methanol as a function of space velocity. Reactant mixture 10% CO/IO% CO/SO% H,, containing 14C0, at 50 bar, 523 K (Chinchen et al., 1987).

interchange. Chinchen et al. (1987) and Chinchen et al. (1990) used carbon-14 labeling of 14C0 or 14C02 to attempt to find the answer. At low residence times, and thus very low conversions, the methanol had the same 14C content as the CO, (Fig. 10). They concluded that the Russian workers were correct; all methanol under these conditions and at higher conversions is made from CO,. They also found that the WGS reaction and the methanol synthesis do not have a common carbon-containing intermediate on the surface of the catalyst. All methanol is made immediately from CO,; CO is converted via the WGS to CO,, which is then converted to methanol. Waugh (1992) has published a review of the results of many years of work at ICI on the kinetics and mechanism of methanol synthesis over Cu/ZnO/Al,O, and other oxide supports. There is little doubt that copper metal is the important catalyst constituent. Waugh and his co-workers concluded that it is the CO, part of the CO, CO, mixture that is the precursor to methanol, being adsorbed on the partially oxidized copper as a symmetrical carbonate which is then hydrogenated to a formate species adsorbed on the copper. The formate is possibly the longest lived intermediate leading to methanol. Hydrogenation of this formate species is probably the rate-determining step in the synthesis. The specific activity of copper in the methanol synthesis is apparently not significantly affected by the nature of the oxide support, so that there may be no unique synergy to the Cu/ZnO combination. The role of CO in the methanol synthesis is probably to maintain copper in a more reduced, hence more active, state than reduction with hydrogen alone could achieve. Nevertheless, we do read that the catalyst of choice is Cu/ZnO/Al,O,. Zinc oxide seems to have a number of functions, perhaps performing somewhat better than other

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21 I

supports such as CrO,, SiO, or MnO. The functions that zinc oxide plays have been listed by Chinchen et al. (1990). Zinc oxide helps give the catalyst a high surface area, especially with Al,O, present, is suitably refractory and hinders the copper particle agglomeration that is inevitable in the life of the catalyst. Zinc oxide also acts as a sink for sulfur and chlorides, which poison copper catalysts. In addition, the oxide interacts with Al,O, in the catalyst to cut down on the conversion of methanol to dimethyl ether. Chinchen and Spencer (1991) have proposed the following scheme for the synthesis of methanol from carbon dioxide and hydrogen over a Cu/ZnO/Al,O, catalyst. Hz(g)

* 2H(a)

(4.8)

CO,(g)

* CO,(a)

CO,(a)

+ H(a) # HCO,(a)

(4.9)

HCO,( a) + 2H( a) e CH,O(a) CH,O(a)

+H(a)

8 CH,OH(g)

(4.10) + O(a)

(4.11) (4.12)

CO(g)

+ O(a) *CO,(g)

(4.13)

Hz(g)

+ O(a) @ H,O(g)

(4.14)

Some of these reactions include several steps. The reaction in Eq. (4.1 l), the hydrogenolysis of adsorbed formate, may be the critical step. The synthesis of methanol from carbon dioxide over Cu/ZrO/Al,O, is an insensitive reaction (Burch et al., 19901, meaning that the specific activity (turnover frequency) under standard conditions depends little on catalyst parameters such as catalyst composition, metal crystallite size, nature of catalyst support, method of catalyst preparation, promoters or poisons (impurities), etc. Sensitive catalytic reactions are strongly dependent on catalyst parameters. Methanol synthesis over Cu/SiO,, for instance, is a sensitive reaction (Nonneman and Ponec, 1990), and this may account for differences in activity of various supported copper based catalysts. In summary, methanol is formed from CO, in the H/CO/CO, feed. Copper is the active catalyst component. The CO, molecule adsorbs dissociatively (CO, + CO f 0) on the copper to form a bidentate formate which is hydrogenated to methanol in a rate-determining step. The role of CO is to keep the copper in a highly reduced state. Oxygen coverage of the copper is a function of the CO/CO, ratio. Although methanol can be made from H/CO feeds on Cu/Zn/Al,O,, the rate is about 100 times slower than when CO, is present. The presence of CO, enhances the durability of the catalyst; in the absence of CO,, the catalyst deactivates more rapidly. Studies devoted to the mechanism of the synthesis of methanol remain under active investigation. The mechanism given above is still under vigorous study. The kinetics of the methanol synthesis are complex and are affected by a number of variables, such as the nature of the catalyst, the physical changes of the catalyst as the reaction progresses, the composition of the gas (which is also constantly changing in the reactor), temperature, and pressure. Modem methanol syntheses use copper-zinc lowpressure catalysts and, as the synthesis reaction proceeds to thermodynamic equilibrium very rapidly, the kinetic behavior of the catalyst may not be of great importance. Commercial catalysts and data on their kinetic behavior are more or less proprietary.

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The production of methanol is an established commercial technology. Nevertheless, constant improvements are being made in process technology and reactor design for better recovery, in lower compression costs, and in processing of the raw methanol. Lee (1990) and Satterfield (1991) have summarized information on these subjects. It is particularly important to avoid contamination of the methanol catalyst by metals that are Fischer-Tropsch catalysts. Care is required in catalyst preparation so as to obtain pure methanol. Nickel and especially iron, both of which form volatile metal carbonyls, Ni(CO), and Fe(CO),, respectively, must not be allowed to come into contact with the syngas under reaction conditions. The carbonyls form at lower temperatures and decompose to the metal and CO at higher temperatures, possibly in upstream heat exchangers, etc. The presence of Fischer-Tropsch metals in the catalyst or on the reactor walls will result in the formation of methane and also of higher hydrocarbons and higher molecular weight oxygenated products. Methanol synthesis reactor shells are typically lined with copper, although internal parts may be constructed of 18-8 stainless steel. The formation of higher alcohols (ethanol, propanol, etc.) can be suppressed by careful exclusion of alkalis from the catalyst. Dimethyl ether is formed by the dehydration of methanol or by the hydrogenation of CO and may form in the presence of Al,O,. However, if a Cu/ZnO low-pressure catalyst is employed with about 7.5% Al,O, as stabilizer and promoter, ether formation is negligible. The copper-zinc catalysts vary in metal composition and contain different amounts of other metals, such as Cr, Al, Mn, V, Ag, etc. Nonneman and Ponec (1990) have shown that a pure copper catalyst is not active for the synthesis of metals; the presence of small amounts of (alkali) promoters is necessary for catalytic activity. The Boudouard reaction, 2C0 @ C + CO,, which results in carbon laydown, is not significant if the temperature is carefully controlled, despite a highly favorable thermodynamic tendency. 4.3. Developments

in methanol synthesis technology

Turnkey plants for the synthesis of methanol are available from a number of companies throughout the world. Newer units for methanol synthesis built by separate manufacturers generally embody what may be termed evolutionary improvements, so that there is a continuous incorporation of new technology. These efforts maintain a well-entrenched viable industrial synthesis of methanol. Completely new routes for the synthesis of methanol are under study but they must show considerable advantages to be considered for commercialization. Meanwhile basic and applied research on the mechanistic aspects of the synthesis of methanol continues to attract the attention of scientists and engineers. Conversion of syngas to methanol in current commercial plants is limited to about 25% by thermodynamic considerations. A higher conversion of syngas per pass can be achieved by operation at lower temperatures to increase allowed levels of methanol. Methanol equilibrium as a function of temperature and pressure is shown in Fig. 11. Efforts are made to keep operating temperatures as low as possible. More active catalytic systems which operate at lower temperatures are being explored, but these are

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6

Technology 48 (1996) 189-297

10

20 Pressure,

Fig.

21.3

30

MPa

I 1.Methanol equilibrium as a function of temperature and pressure.

generally deactivated by low levels of CO, (< 100 ppm). The expense of removing CO, from syngas hampers the adoption of these routes for methanol synthesis. The thermodynamic constraint to methanol synthesis can be overcome by shifting the equilibrium by removing methanol during reaction so as to attain higher conversions per pass. This promising approach may well be incorporated into newer commercial units. Reviews of newer catalytic systems have been published by Mills (1988, 1993) and by Trimm and Wainwright (1990). 4.3.1. Alloys as catalysts Baglin et al. (1981) discovered that Cu/ThO, catalysts prepared from intermetallic CuTh, alloys were active in methanol synthesis. This system was also studied by Daly (1984). Owen et al. (1987) probed into a larger class of intermetallic CUM,~ (M is Ce, La, Pr, Nd, Gd, Dy, Zr, Ti, or Th) alloys as precursors of Cu/M,O, and Cu/MO, catalysts, and found that the CuCe,,, CuLa,,,,, and CuPr,,, precursors also resulted in highly active catalysts for the synthesis of methanol. Methanol synthesis using Cu/La catalysts was achieved at temperatures as low as 100°C. These catalysts were severely poisoned by very small amounts of CO, and often contained copper metal of low dispersion (very large particles, Cu metal area < 1 m2 g- ‘1. An electropositive copper species analogous to that proposed by Klier et al. (1986) for the Cu/ZnO catalyst, by Baglin et al. (1981) for the Cu/ThO, catalyst, and by Shibata et al. (1984) for the Cu/ZrO, catalyst was tentatively suggested to be the active component. Alternative candidates for the site for methanol synthesis were suggested to be extremely small copper particles ( < 1 nm diameter) or inter-metallic hydrides. Because of the lack of tolerance to CO,, the Cu intermetallics-derived catalysts. although very highly active, are not considered practical, as industrial syngas invariably contains significant amounts of CO,, and CO, removal to very low levels adversely affects the process economics. The oxygenate selectivities of these catalysts of 80-98% are acceptable in their upper limit but not in the lower limit. Taking into account their very high activities, the lack of precise determination of CO, effects at low concentrations, and the lack of data on water effects on the synthesis, it seems that these novel copper based catalysts have not been studied in detail and further research into their possible improvement is warranted.

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4.3.2. Cs / Cu / ZnO catalysts Cu/ZnO catalysts can be doped with the more active alkalis to increase the yield of methanol from syngas low in CO, (Klier et al., 1986; Nunan et al., 1988; Klier et al., 1982). Catalytic activity depends on the level of alkali doping of Cs in the Cu/ZnO = 30/70 mol% catalyst and the ternary Cu/Zn/Cr = 30/45/25 mol% catalyst. Methanol selectivity at 250°C and 7.6 MPa was over 97% with H.&O = 2.33 at optimum Cs levels. Side products were mostly ethanol and methyl formate. The high selectivity to methanol can be shifted towards higher oxygenated compounds by changing the reaction conditions. The use of potassium instead of cesium gives poorer results. 4.3.3. Supported Pd catalysts Poutsma et al. (1978) showed that methanol could be formed selectively over Pd/SiO, catalysts with a 70:30 ratio of H,/CO at 260-350°C 5-110 MPa and GHSV = 3300 h- ’. A number of different supports and alkali additives have been studied. van der Lee et al. (1986) have discussed this somewhat surprising methanol selectivity of Pd based catalysts in view of the tendency of Group VIII transition metals to function as IT or methanation catalysts. Although Pd catalysts can selectively yield oxygenates from H,/CO = l-3 syngas under conditions of commercial interest, yields of methanol are low and industrial incentive to pursue their use is lacking. 4.3.4. Brookhaven liquid low-temperature and related syntheses

This methanol synthesis was originally based on NaI-I/RONa/M(OAc), catalysts at the Brookhaven National Laboratory (Slegier et al., 1984; Sapienza et al., 1986). The catalyst is completely dissolved in a solvent medium, unlike heterogeneous catalyst systems. It is probable that the later catalytic systems are composed of an alkali methoxide such as KOCH, and a nickel salt in a solvent such as triglyme. The synthesis takes place at 80-12O”C, yielding > 90% syngas conversion per pass with over 95% selectivity to methanol. A selectivity to methyl formate of up to 78 mol% can be achieved in the same system (Mahajan and Mattas, 1992). With nickel as part of the catalyst, the formation of the extremely toxic nickel carbonyl must be guarded against. The process has been operated using syngas made by partial oxidation using air to avoid the expense of an oxygen plant. The system is sensitive to small amounts of CO, which, as stated, is expensive to remove from syngas to low levels. However, this process is a significant departure from the commercial synthesis of methanol (Trimm and Wainwright, 1990). Marchiomra et al. (Marchiomra et al., 1988, Marchionna et al., 1992) have investigated systems which produce methanol from syngas at similar low temperatures and pressures (80-120°C lo-50 atm). The active catalytic system is again homogeneous and involves a Ni(CO),/NaOCH 3 combination. Infrared investigations indicated that attack by CH,O- on Ni(CO), under a hydrogen atmosphere produces (HNi(CO),)- and methyl formate through the intermediate formation of (Ni(CO>,(COOCH,)-. Methyl formate is also synthesized by the NaOCH,-catalyzed carbonylation of methanol. Then the anionic hydronickel carbonyl species catalyzes the reduction of methyl formate to methanol via the intermediate formation of formaldehyde (HCHO). The extremely mild methanol synthesis conditions were postulated as being due

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215

chiefly to the more facile reduction of the activated carbon-oxygen bond in methyl formate as compared with that of CO coordinated to a metal. This low-temperature methanol synthesis resembles that of the Brookhaven synthesis. 4.3.5. Concurrent methyl alcohol/methyl formate synthesis Christiansen (1919) patented a two-step synthesis of methyl alcohol via the following reactions CH,OH + CO * NaOCH, + HCOOCH,

(4.15)

HCOOCH, + 2H, P 2CH,OH

(4.16)

Net 2H, + CO @ CH,OH

(4.17)

The first step is catalyzed by an alkali alkoxide (NaOCH,, KOCH,, etc.); the hydrogenolysis reaction in the second step is carried out in the liquid phase using a catalyst such as copper chromite. Both reactions in Eqs. (4.15) and (4.16) have been investigated in some detail (Tonner et al., 1983). The alkoxide catalysts react with water and with CO, to form alkali formates (e.g. HCOOK) and alkali methyl carbonates (e.g. KOCOOCH,), respectively (Liu et al., 1988, Liu et al., 1989) so that great care must be taken to remove water and CO, to levels of about 1 ppm and 10 ppm, respectively. These deactivating effects have deterred investigations in which both steps, Eqs. (4.15) and (4.161, would be carried out in a single reactor. However, Liu et al. (19881, Onsager (1984) and Aker Engineering (1982) have reported the synthesis of methanol directly from syngas in a single reactor using a catalyst comprised of an alkali methoxide and copper chromite; this has been referred to as the concurrent synthesis of methanol. Under the conditions of this reaction, 1OO- 180°C and 50-65 atm, there is an interaction between the carbonylation and hydrogenolysis catalysts, with the overall rate being higher than that predicted for either reaction at these conditions. The copper chromite catalyst regenerates the carbonylation catalyst which has been deactivated by water and also removes water by the WGS reaction (Palekar et al., 1993a). Indeed, potassium formate and potassium hydroxide used with copper chromite are as active as the potassium methoxide/copper chromite mixed catalyst in this concurrent synthesis (Palekar et al., 1993b). The reaction is carried out in a slurry with consequent good heat transfer rates and with a rate of reaction at low catalyst loadings which is comparable with that in the commercial synthesis. High per pass conversions can be obtained, and the only significant products are methanol and methyl formate, which can be separated easily. However, deactivation by water and CO, is severe. 4.3.6. The liquid-phase methanol synthesis Sherwin and Frank (1976) at Chem Systems, Inc., developed a liquid-phase methanol synthesis (LPMEOH’“). This slurry-phase process has a number of advantages: it can use syngas rich in CO as obtained from modem coal gasifiers, enhanced heat transfer of the highly exothermic heat of reaction, and a high conversion per pass. Researchers at the University of Akron have been investigating various aspects of this process in a 1 liter slurry reactor that closely approximates CSTR conditions (Lee, 1990;

216

I. Wender / Fuel Processing Technology 48 (1996) 189-297

Lee et al., 1992). Using a Cu/ZnO/Al,O, catalyst, they established reaction chemistry for diverse gas feeds including syngas with high H,, high CO, CO,-free and CO-free environments. Their studies led to the suggestion of a liquid entrained reactor with a catalyst slurry for commercial operation of the liquid-phase methanol synthesis. In 198 1, the US Department of Energy began supporting research on a liquid-phase methanol synthesis in a process development unit at the LaPorte, TX plant operated by Air Products and Chemicals, Inc. (Brown et al., 1991). The Electric Power Research Institute co-sponsored the program because of their interest in the possible coproduction of methanol and electricity in integrated gasification combined cycle (IGCC) systems. Operational conditions for methanol synthesis were studied in a 10 ton per day unit with a Cu/ZnO/Al,O, catalyst. In situ catalyst activation was achieved with a l-10 pm sized catalyst in Witco 70 oil with reaction temperatures of 225-265°C (Lewnard et al., 1990). An extensive series of tests was run with various CO/H, ratios. Catalyst slurries of 20-45 wt% were used at 250°C with CO-rich syngas. Methanol productivity was optimum when the CO, content in the syngas with a H,/CO ratio of 0.69 was 5-8 mol% (Herman, 1991). This work demonstrated the viability of the slurry bed concept, attaining a CO conversion to methanol of about 13% per reactor per pass in a hydrogen-rich feed. A primary output of the Eastman Chemical Company plant in Kingsport, TN is carbonylation-derived acetic anhydride (see Section 7). At present, methanol for acetic anhydride is produced from syngas by a gas phase process. The Air Products’ liquid-phase methanol synthesis enables sufficient heat removal without the need for a a WGS reactor, as mentioned earlier. Construction of a demonstration plant at Kingsport, with partial funding by the US Department of Energy, is underway in a joint Air Products/Eastman venture (Cook, 1995). Start-up of this plant may take place in 1997. 4.3.7. Methanol removal to obtain higher conversions The synthesis of methanol is a classic example of an exothermic equilibrium-limited reaction. There is recent emphasis on improving the synthesis of methanol by removing the alcohol as it is formed, so lowering the thermodynamic constraints on methanol conversion. Berty et al. (1990) introduced the concept of “beating the equilibrium” by introducing a high-boiling inert solvent such as tetraethylene glycol dimethyl ether (tetraglyme) into a reactor containing a fixed catalyst bed. The solvent, flowing concurrently with the syngas stream, absorbs methanol as it is formed. This results in an equilibrium shift to methanol, as the activity of methanol in the reactor remains low. This concept, based on the classical idea that the equilibrium limitation in a reversible reaction can be overcome by product removal on its formation, works well; syngas conversion is enhanced and recycle is almost eliminated. The principle is sound but has not so far been used in a commercial synthesis of methanol. Westerterp and Kuczynski (19861, using a gas-solid trickle flow reactor (GSTPR), have obtained methanol yields of 100% by removing methanol by selective adsorption on an amorphous SiO,/Al,O, cracking catalyst in the reaction zone. The same group (Westerterp et al., 1988; Westerterp et al., 1989) has employed another approach to the separation of product methanol by use of a reactor system with interstage removal of the

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alcohol between two packed tubular reactors in series (RSIPR). Methanol is absorbed at reaction temperature in a countercurrently operated packed bed absorber using tetraglyme (boiling point 275°C) as the solvent. Here a liquid rather than a solid is used to absorb product methanol. The authors envisage significant energy and raw material savings on scale-up of this system. 4.3.8. Condensing methanol principles Hansen and Joensen (1991) at Haldor Topside A/S have carried out a study to ascertain the best conversion to methanol obtainable per pass with a 2:l Hz/CO gas mixture and with an acceptable space time yield. This approach is essentially based on the finding that maximum conversion obtainable for a stoichiometric gas feed for the synthesis of methanol decreases as the CO, content increases. A distinct carbon conversion maximum is observed at about 2% CO, in the make-up gas, in line with the findings of Klier et al. (1982). These workers generated a syngas with about 67% H,, 29% CO and 3% CO,. Conversion to methanol was calculated assuming that both the synthesis of methanol and the WGS were in equilibrium. Calculated conversion corresponded to such high partial pressures of methanol that condensation of the methanol and water would occur. The calculated dew point lines as a function of temperature, conversion level and pressure are depicted in Fig. 12, as are the equilibrium curves. Condensation takes place above and to the left of the dew point line. Proximity to the critical point of methanol (239.43”C, 79.9 atm) is responsible for deviation from an essentially straight dew point line for 12.5 MPa at high conversion. Results of laboratory experiments with conversions at 230°C and 240°C carried out in isothermal reactors agreed well with those predicted for gas-phase equilibria (Fig. 13). At 220°C and below, levels of conversion were high enough so that condensation occurred and the dew point line was crossed. At 200-22O”C, conversions were higher than those predicted by gas-phase thermodynamics. These findings are explainable only by the highly non-ideal properties of the liquid phase on the catalyst, essentially by the fact that the activity coefficients in the liquid phase deviate to a large extent from unity.

Fio 12. Methanol + shift equilibria fr& Hansen and Joensen. 1991).

and dew points. (Reprinted

with permission

of Elsevier Science Publishers

218

I. Wender/

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Technology 48 (1996) 189-297

s m lm 356

.Bfeamld

lso 374

an 392

zlo 410

z iii

24072 44°F

Fig. 13. Once-through conversion of CO and CO, 9.6 MPa; CO = 30%. (Reprinted Elsevier Science Publishers from Hansen and Joensen, 1991).

with permission

of

Haldor Tops& carried out pilot plant work and concluded that their methanol synthesis with a newly developed catalyst, combined with two-step reforming (a smaller first-stage steam reforming followed by an autothermal reforming (Hydrocarbon Processing, 1994) could be an economic way of producing methanol. 4.3.9. Methanol / dimethyl ether from syngas The physical removal of methanol is one way of overcoming thermodynamic limitations in the synthesis of methanol. Another interesting and promising way is to convert methanol into a chemical species whose removal affects the equilibrium conditions. The in situ dehydration of methanol to dimethyl ether (DME), in the presence of an added acid catalyst such as gamma-Al,O,, is based on the second option. The direct synthesis of DME from synthesis gas by dual function catalysts has been investigated by a number of companies (Mobil Oil, 1975; Slaugh, 1983; Haldor Top&, 1985; Brown et al., 1991). The synthesis of DME in a single reactor is based on a combination of an equilibrium-limited reaction (synthesis of methanol) and an equilibrium-unlimited reaction (methanol dehydration over the added acid catalyst) (Lee et al., 1992). The reactions involve methanol synthesis, methanol dehydration, and the WGS reaction 4H, + 2C0 e 2CH,OH AH = -43.2kcal

(4.18)

2CH,OH @ CH,OCH, + H,O AH = -5.6kcal

(4.19)

AH= -9.8kcalmol-’

(4.20)

CO+H,0+H2+COZ

Net 3H, + 3C0 F? CH,OCH, + CO, AH = -58.6kcal

(4.21)

Here one of the products in each step is a reactant for the other; hydrogen formed as in Eq. (4.20) is a reactant for the methanol synthesis. This is a strong driving force for allowing very high conversions of syngas for the overall reaction. The combination of reactions results in a synergistic effect that eases the thermodynamic constraints for the

1. Wender/

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48 (1996) 189-297

219

(CO + CO 2 >

% Conversion

MeOH + DM3

MeOH

0

1

2

3

Pressure

4

5

6

7

6

@IPa)

Fig. 14. Equilibrium conversion of syngas versus pressure Science Publishers from Hansen and Joensen, 1991).

at 240°C. (Reprinted

with permission

of Elsevier

methanol synthesis. The data in Fig. 14 show the advantage of combining the methanol and DME syntheses. Conversion levels similar to those achieved in conventional methanol synthesis may be attained at much lower pressures. The process is thus an attractive front end for those syntheses where DME may be employed as a feedstock instead of or with methanol (Hansen and Joensen, 1991). DME has potential for use as a motor fuel, for energy in the TIGAS process and as an storage, as a propellant in aerosol containers, intermediate in the synthesis of chemicals (Rostrop-Nielsen, 1994). Brown et al. (1991) compared the results of the liquid-phase DME (LPDME) system with those of the LPMEOH process using a number of methanol and dehydration catalysts. They found that the conversion of CO in the LPDME system was almost twice as high as that achieved with LPMEOH and much higher than the conversion obtained in the reaction of methanol alone. Work by Brown et al. (1991) and also the contribution of Gogate et al. (1993) have shown that the single-step DME synthesis from coal-derived syngas gives greater syngas conversions than that achievable in the LPMEOH process. The route to methanol synthesis via DME has great potential for enhanced overall process performance. Future work will be necessary to determine the effects of process variables such as temperature, pressure and feed gas composition and also the activity and stability of the mixed methanol/dehydration catalyst system. The reaction of the methanol/DME system provides a challenge to conventional methanol manufacturers to design more active systems for the synthesis of methanol. 4.4. Fuel uses of methanol Methanol is a fuel and a chemical and is used in the synthesis of a wide variety of fuels and chemicals. As mentioned earlier, fuels are sold by the ton, chemicals by the pound. The huge amount of methanol consumed worldwide per year (24 million tons and still growing) can only mean that the main growth outlets for the uses of methanol must be as a fuel, essentially in liquid transportation fuels. The chief uses of methanol for chemicals will be discussed in a later section of this review.

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The main uses of methanol as a fuel, essentially in transportation fuels, are listed below. * Neat methanol or in mixture with gasoline in vehicles-commercial - Methyl ethers as octane enhancers and fuels-commercial - Higher alcohols (C ,X6) as octane enhancers and fuels-not now commercial 4.4.1. Methanol, neat or in mixture with gasoline In the limited time of high gasoline prices, there was a strong incentive to use methanol or methanol-containing products in gasoline. The ban on the use of tetraethyllead and the Clean Air Act Amendments (CAAA) of 1990, which advocate a lowering of aromatics in gasoline, resulted in a turn to the use of octane-enhancing oxygenates in gasoline. It appeared that a new potential market for methanol was opening up as a motor fuel; emissions of CO, NO, and hydrocarbons are claimed to be low, and it has octane-enhancing properties. A review of research on and the industrial activity of methanol has been published by Sinor Consultants (1990). It is well to pause here to consider that the enactment of the CAAA in 1990 is somewhat revolutionary in scope. The United States government, for the first time, is in a sense regulating the composition and properties of gasoline and the kinds of transportation fuels sold as these concern environmental and health policy (Peeples, 1991). Provisions of the Act require manufacture and sale of clean fuels which will reduce refueling, evaporative and exhaust emissions of ozone-forming compounds, airborne toxic compounds and CO from motor vehicles. The petroleum industry and car manufacturers have cooperated in meeting the challenges raised by the CAAA. ARC0 led the way, followed quickly by other major manufacturers of motor fuels. Refinery blending of methyl ethers such as methyl t-butyl ether (MTBE) helped meet the regulated oxygen requirements, allowed the removal of high-octane aromatics and helped lower or balance the vapor pressure and the light olefin content of gasoline. The oxygenates that are most likely to be included in gasolines in the United States are shown in Table 1 (Unzelman, 1992). As will be mentioned later, alcohols with four to six carbon atoms blend well with gasoline and can be added directly to gasoline. Table 1 Oxvgenates

most likelv to be used in United States gasoline Blend octane, (R + M)/2

Blend RVP/psi

Blend b.p./“F

Oxygen/W%

Alcohols TBA Ethanol Methanol

101 113 116

10-15 17-22 50-60

181 172 149

49.9 34.7 49.9

Ethers MTBE TAME ETBE

109 104.5 110

8-10 3-5 3-5

131 187 161

18.2 15.7 15.7

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Methanol, however, has a number of drawbacks; the most serious is phase separation, which would create serious problems in transportation, distribution and use in mixture with gasoline. With pure methanol, 500 ppm of water is enough for two phases to form. Corrosion caused by methanol-gasoline blends is also a problem. In the early 198Os, ARC0 sold a methanol plus cosolvent mixture with gasoline in parts of New York State and Pennsylvania under the trade name “Oxinol”. It consisted of 4.75% of methanol and 4.75% of t-butyl alcohol (TBA) in unleaded gasoline. Meanwhile Midwestern refiners sold ethanol-blended gasoline. A drop in demand for Oxinol resulted, mainly through lack of consumer acceptance of methanol in gasoline and higher methanol prices; it is no longer in use. Racing cars use methanol with 15 ~01% of gasoline (M-8.5); gasoline is needed because methanol burns without a visible flame. Neat methanol has been tested in specially adapted cars in California and other places but distribution and costs, among other problems, appear to hinder this approach from making a significant contribution to our transportation fuel needs. 4.4.2. Methyl ethers and higher alcohols as octane enhancers and fuels To overcome air pollution in certain cities in the United States, the concept of reformulated fuels in automobiles has been introduced. Gasoline-fueled vehicles emit small amounts of oxides of nitrogen (NO,), CO, hydrocarbons and fuel from tanks and fuel-delivery systems (Seddon, 1992). By the mid- 1980s the concept of the use of alternative, clean-burning fuels had gained increasing support. United States oil refiners introduced “reformulated gasoline”, a concept embraced by the Clean Air Act of 1990. According to this Act, 38 US cities with high levels of CO would have to use reformulated gasoline (RFG) with an average oxygen content of 2.7 wt%. Gasoline in cities with high levels of ozone would have to reduce volatile olefins by 15% by 1995 compared with 1990 levels. Oxygen content in cities with high ozone levels would have to be an average of 2 wt%. The CAAA legislation resulted in the introduction of alternative fuels not derived from petroleum. Alternative fuels could be mixtures comprised of 100% gasoline, with a composition primarily limited to a balanced mixture of normal butane, isoparaffins and toluene to sustain octane ratings with low photochemical reactivity and low emissions. A very large overhaul of refining operations would be required to achieve this goal. Refiners responded with the introduction of RFG which would have performance equivalent to that of alternative fuels, so avoiding the enormous capital expenditure and massive write-off on fuel supplies and distribution (Seddon, 1992). An illustration of RFG proposed compositions which pre-date the CAAA provisions is given in Table 2 (Seddon, 1992). They were formulated with the purpose of lowering levels of CO, which has a propensity to form ozone; an average of 2.7 wt% of RPG would be required in cities with high CO levels. Volatile olefinic compounds (VOCs) would be eliminated by 1995 compared with 1990 levels and benzene would be limited to 1 ~01%. Total aromatics would be limited to 25 ~01%. The data in this table depict the refiners’ efforts towards the marketing of oxygenates in gasoline with limited photochemically active compounds. All of the reformulated gasolines in Table 2 contain the compound methyl t-butyl

I. Wender / Fuel Processing Technology 48 (1996) 189-297

222

Table 2 Some previously proposed reformulated unleaded gasolines (Seddon, 1992, reprinted with permission from Elsevier Science Publishers) Company

Arco

Phillips

Conoco

Shell

Exxon

Product

EC- 1

Unleaded plus

RXL

SU2OOOE

Supreme plus

Supreme

87 I .o-2.0 MTBE < 20 < 10

86.5 0.0-2.5 MTBE < 25 no limit < 2.4 8.5

91 1 MTBE no limit no limit no limit 8.0-8.5

89/91 1 MTBE no limit no limit no limit 8.5

91 1 MTBE no limit no limit no limit reduced

Octane, (R + M)/2 a Oxygen b/wt% Oxygenate Aromatics/vol% Olefms/vol% Benzene/wt% RVP/psig

88

I MTBE c 20 < 10 < 1.2 8

I 8.5

Chevron

a (RON + MoN/2. b Lower value for summer, higher for winter.

ether (MTBE), the most common oxygenated compound in use in gasolines in the world. In the period 1982-1992, MTBE production increased at an annual rate of 30%. MTBE production, expected to reach 31 million metric tons by 1995, will then probably level off into the late 1990s. MTBE is produced by the acid-catalyzed reaction of methanol and isobutene; methanol contributes 35.3 wt% of the ether (Eq. (4.22)). Mixed olefin streams from FCC units are used in the manufacture of MTBE. There have been concerns about obtaining an adequate supply of isobutene but ongoing research promises to solve this problem. MTBE is an excellent octane booster, has little effect on volatility, helps reduce CO emissions and has no adverse effect on vehicle system materials. It has favorable physicochemical properties compared with alcohols, including a higher energy content than methanol or ethanol. CH,OH + (CH,),C= isobutylene

CH$CH,

- 0 - C(CH,),

(4.22)

MTBE

There is a growing interest in t-amyl methyl ether (TAME), which uses the C, fraction from FCC units and from steam crackers (ethylene plants). It is estimated that TAME could supply about 70 000-90000 bpd of oxygenates in gasoline before 2000 (Unzelman, 19921, perhaps eventually equalling about a tenth of total ethers. ETBE (ethyl tertiary butyl ether) and TAEE (tertiary amyl ethyl ether) can play roles similar to that of MTBE. Indeed, they have the advantage of lower blending vapor pressure (a lower RVP is of great importance) and higher boiling points. At present, fuel ethanol is not produced from syngas but rather by fermentation of starch. Gasohol, a 9O:lO mixture of gasoline and ethanol, now constitutes about 1% of the gasoline pool. The manufacture of ethanol by fermentation carries a federal and often a state subsidy. To a large extent, then, the continuing growth in methanol synthesis is tied to the increasing use of methyl ethers in reformulated fuels. The synthesis of mixtures of methanol and higher alcohols from syngas by modification of reactions in the synthesis of methanol and by alkali promotion of methanol

5

4

(Japan)

C, Chemistry

Group

Carbide)

Dow Chemicals

(Union

Rh, Mn + Cu, Zn

MoCoS

Li

K

K

Al, Cr, Zn,

H*S

others

Ti, Mn, Zn,

CuNi

NaK

Idemitsu-Kosan

cuco

No

others

IFP

K, others

La, VCe

Mn, AI MO, Ti

Others

K

metals

Na

Alkali

’

I500

30000-45000

5000-7000

300%6OQO

2OOt-4000

3oOt-

GHSV/h-

260-280

5

12-14

6-10

260-320

290-310

7-10

(18)

12-16

P /MPa

270-300

350-420

T,“C

conditions

et al., 1990)

Range of operating

projects (Courty

and

3

CuAlZn

CU

Haldor Topwe

Lurgi

Zn

Cr

Snamprogetti

Key elements

The catalysts

status in 1988 for Syngas to alcohol

Enichem

2

I

Company

Research and development

Table 3

I .4

1.1-1.2

l-2

l-I.2

OS-3

H, /CO

removal

No

Yes

(synth.

loop)

Yes (OS-3%)

Yes (1% synth. loop)

No

CO,

Demonstration

Lab. scale

(28 I cat.)

Bench scale

(20 bbl/dayJ

unit

Italy)

plant t/y,

Bench scale

(I5000

Industrial

State of development

.4

% !!.I s

2 ;= z 8

2 $ 0 5 B

1 < 2 x $ E’ h

7

k

f

224

I.

Weder/

Fuel Processing Technology 48 (1996) 189-297

synthesis catalysts has been known since 1913 (Mittasch and Schneider, 1913). Indeed, from 1927 to 1945, mixed alcohols were manufactured in Germany using alkalized iron catalysts. Early work on the higher alcohol synthesis (HAS) has been reviewed by Natta et al. (19571, Anderson et al. (1952) and Stiles (1977). In the 1980s reviews of the HAS were published by Haag et al. (1987), Xiaogding et al. (1987), Mills (1988) and Wender and Klier (1989). More recently, excellent reviews of the HAS have been published by Courty et al. (19901, Herman (1991), Forzatti et al. (1991) and Mills (1993). The HAS consists of the conversion of syngas to methanol plus higher alcohols; methanol is usually the chief component of these higher alcohols (Smith and Anderson, 1983; Stiles et al., 1991). It is pertinent to realize that, although there is growing use of the addition of MTBE and other methyl ethers to gasoline in reformulated fuels, the Snamprogetti/Enichem, Tops&e, Lurgi, Dow/Union Carbide, Institut FranGais du Petrole, and Vulcan projects on the HAS were all abandoned before a single commercial plant was built (Notari, 1991). The status of research and development in 1988 of the various projects for the synthesis of higher alcohols is given in Table 3. Mixed alcohols are inferior as blending agents in gasoline and, although their use is not precluded, they are not presently regarded as commercially promising. The composition of some of the higher alcohols prepared from syngas is given in Table 4 (Mills, 1993). The presence of lower alcohols in the HAS mitigates against their use. The large amounts of methanol, in particular, in the C,-C6 alcohols increases their susceptibility to extract out of gasoline under wet storage. Ethanol, for example, is not usually blended with gasoline at the refinery but is added to gasoline just before delivery to the gasoline pump (Piel, 1993). El Sawy (1990) has published an evaluation of mixed alcohol production catalysts and processes. It is difficult to rank the various HAS processes because of their different responses to areas of concern for the EPA. These include responses and activity in relation to exhaust emissions, evaporative emissions, driveability, materials compatibility and fuel stability (phase separation). Methyl ethers such as MTBE and, to a lesser but growing extent, TAME and especially ETBE are regarded as having much better properties as gasoline blending agents. The ethers are more effective and have less costly pathways to deliver alcohols and energy from domestic sources such as natural gas or coal into the liquid transportation fuel system of the United States (Piel, 1993). Table 4 Composition of some fuel alcohols from syngas Alcohol/%

C,

C,

C,

C,

C,

Catalyst

MAS (SEHT) Substifuel (IFP) Octamix (Lurgi) HAS (Dow)

69 64 62 26 a

3 25 7 48

4 6 4 14

13 2 8 3.5

9 2.5 19 0.5

K/Zn/Cr K/&/Co/Al akali/Cu/Zn/Cr CoS/MoS, /K

a Methanol can be recycled to extinction, increasing the amount of ethanol. Dow: straight chain alcohols. Lurgi: isobutanol is 70% of C, alcohols.

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225

Higher alcohols, with four to six carbon atoms, are good gasoline blending agents and may be added directly to gasoline. They are essentially as good as methyl ethers in blends with gasoline. Isobutanol, t-butanol (TBA), pentanols and hexanols may be used in reformulated fuels as such. They have low water solubihty, low RVP properties, high energy contents and are completely soluble in gasoline. The potential octane contribution of oxygenates in gasoline as a function of their oxygen contribution in gasoline is illustrated in Fig. 15 (Pie], 1994). Alcohols are currently limited to 3.7 wt% oxygen (methanol to 2.5 wt%); ethers are limited to 2.7 wt%. With these limits, ethers and alcohols could replace from 1.5 to 20% of gasoline. The replacement potential as energy is shown in Fig. 16 (Piel, 1993). The energy replacement potential of the oxygenates is slightly lower than the volume replacement. The C, alcohol isobutanol is of particular interest. It can be synthesized in fairly good selectivity directly from syngas, is a good blending agent with gasoline, and can be used in the manufacture of MTBE. There is a large amount of isobutanol in the C, fraction of mixed alcohols; as much as 70% has been reported by Lurgi (Mills, 1993). But the selectivity in the synthesis of isobutanol from syngas is not high in the HAS. Thermodynamically, the amount of isobutanol from syngas can be high under certain conditions; it is the dominant alcohol up to 500°C (Roberts et al., 1992). The concept of reacting methanol and isobutanol to form methyl isobutyl ether (MTBE) is attractive; unfortunately, methyl isobutyl ether (MIBE) in unleaded gasoline has the very low BRON (blending research octane number) of 60.8 and a BMON (blending motor octane number) of 67.0 (Herman et al., 1994). The corresponding values of MTBE are 120.1 and 96.3 respectively.

10 IPTBE

:

.’ :

:

ETBE

: .’

:

MTBE

: Ethanol

T-Butanol

Current Oxygen Limit 2

3

4

5

6

Wt % of Oxygen in Gasoline Fig. 15. Potential octane contribution of oxygenates in gasoline (Piel, ETBE, ethyl t-butyl ether. MTBE,

1994).

IPTBE,

methyl t-butyl ether; t-butanol, t-butyl alcohol.

isopropyl butyl ether;

1. Wender / Fuel Processing

226

Technology 48 (1996) 189-297

35 ??

Current Oxygen Level

IPTBE :

30 -

: : : .’ : ,’ : : : : ,’

$

g

2s

Lu _m E

2

::

20_

.’ : .’ :

0

1

2

3

:

ETBE MTBE TAA

,’.’

4

TBA

5

6

7

Wt % Oxygen in Gasoline

Fig. 16. Potential volume contributions (Piel, 1993).

of oxygenates

in gasoline.

TAA, t-amyl ether, IBA, isobutyl

alcohol

The market demand for oxygenated fuels shows a definite trend to higher molecular weight ethers and the more highly branched higher alcohols which have a high energy density, high octane rating, low RVP and high tolerance for water. Typical properties of fuel alcohols and fuel ethers in gasoline are given in Table 5 (Piel, 1994). The gasoline of the future is still of uncertain composition. Very likely, a number of combinations of hydrocarbons and oxygenates will be used, but the situation is fluid at this time. There will be many formulations proposed and used, for the transportation fuel situation is still undergoing rapid change. 4.5. The Fischer-Tropsch

and related reactions

4.5. I. Background and scope The Fischer-Tropsch process was the first one to be used to convert syngas to liquid

fuels on a large scale. The two major reactions involving the hydrogenation of CO to liquids are the Fischer-Tropsch synthesis (FTS) and the synthesis of methanol. The two reactions differ in a major fundamental way: CO is adsorbed dissociatively, to a great extent, in the ITS; in the synthesis of methanol, CO is exclusively adsorbed associatively (without splitting of the C-O bond). The seminal paper in regard to the catalytic hydrogenation of carbon monoxide (to methane) was published by Sabatier and Senderens (1902). In 1913, BASF obtained patents on the preparation of hydrocarbons and related oxygenated compounds by

I. Wender/Fuel Table 5 Typical properties

of fuel alcohols

Technology 48 (1996)

189-297

227

in gasoline (Piel, 1994)

Octanes: blending (R + Ml/2 Vapor pressure near RVP ( lOOoF) Blending RVP b Boiling pt./OF Density/(lb gal- ’) Energy density/(MBTU gal- ’) (LHV) Heat of vapor/(MBTU gal-‘) at NBP Oxygen content/wt% Sohtbility in water/wt% Typical properties

Processing

Methanol a and co-solvent

Ethanol

Isopropanol

t-Butanol

Isobutanol

t-Amy1 alcohol

108+ 4.6 31f 148 6.63 56.8 3.14 50

Il.5 2.3 18 173 6.61 76.0 2.39 34.8 I

106 1.8 14 130 6.57 87.4 1.90 26.7 I

100 1.7 9 181 6.59 94.1

I

102 0.6 5 226 6.71 95.1 1.67 21.6 10.0

97 0.7 6 216 6.79 100.1 1.58 18.2 11.5

I

of fuel ethers in gasoline MTBE

Octanes: blending (R + M)/2 Vapor pressure near RVP ( 100°F) blending RVP Boiling pt./OF Density/(lb gal- ’) Energy density/(MBTU gal- ‘) (LHV) Heat of vapor/(MBTU gal- ’) at NBP Oxygen content/wt% Solubility in water/wt%

1.55 21.6

ETBE

DIPE

TAME

IPTBE

AEE

110

112

105

105

113

100

7.8 8 131 6.19 93.5 0.86 18.2 4.3

4.0 4 161 6.20 96.9 0.83 15.7 2.6

4.9 5 155 6.10 100 0.90 15.7 2.0

2.5 2.5 187 6.41 100.6 0.90 15.7 2.0

2.5 2.5 188 6.30 NA NA 13.8 NA

1.2 1 214 6.39 NA NA 13.8 NA

I: Infinite solubility; NA: not available. ’ Typical for methanol waivered blends with cosolvents. b Blending RVP for 2.7% oxygen or higher in gasoline.

hydrogenation of CO at high pressure, mostly on oxide catalysts. In 1923, Fischer and Tropsch obtained a large amount of oxygenated products from syngas using alkalized iron and other catalysts; they called the product “Synthol”. In the same year, Fischer and Tropsch synthesized higher hydrocarbons using nickel and cobalt at one atmosphere (BASF had patents covering higher pressures). Several countries, including the United States, England and Japan, initiated studies on the FTS starting about 1926. Fischer and his co-workers at the Kaiser Wilhelm Institute for Coal Research (KWIK-now the Max Planck Institute for Coal Research at Mtilheim) developed Ni-ThO,-kieselguhr and Co-ThO,-kieselguhr catalysts for the FTS. Nickel catalysts gave too high yields of methane and work with this catalyst was discontinued (Anderson, 1984). In the decade following 1935, the FTS was operated in Germany using cobalt catalysts, usually in the 0.5-2.0 MPa range (so-called medium pressure). Cobalt was expensive and in short supply; after World War II cobalt catalysts were replaced with alkalized iron catalysts. During this period, Pichler (1952) produced high molecular weight hydrocarbon waxes using ruthenium catalysts at high pressures. Fischer and Pichler, while investigating new catalysts and studying the mechanism of the FTS, developed a process called the Isosynthesis (Cohn, 1956). Roelen (1943), in the course

228

I. Wender/

Fuel Processing

Technology 48 (1996) 185-297

of studies on the mechanism of the ITS, discovered the homogeneously catalyzed hydroformylation (0x0) reaction, now used extensively around the world to produce aldehydes and alcohols from olefins and syngas (this reaction is discussed in Section 5). The production of ITS products in Germany during WWII reached a maximum of about 650000 tons per year in 1944. The standard catalyst used in all the plants had the composition in relative mass units of Co:ThO,:MgO:kieselguhr = 100:5:8:200. Bomb attacks reduced the output of these plants to about 10000 tons per year by early 1945. All these FTS plants had ceased operation by the end of the war. After the war, Ruhrchemie and Lurgi jointly developed a process based on the use, at medium pressure, of iron catalysts in fixed bed units which was commercialized at Sasol 1 and later referred to as the Arge process. In 1948, Koelbel and Englehardt developed a hydrocarbon synthesis from CO and steam to utilize the then available CO-rich blast furnace gas (Koelbel and Englehardt, 1951; Koelbel, 1957). In 1935, the British Fuel Research Station initiated work on the FYI’Susing both cobalt and iron catalysts in fixed bed reactors (Hall et al., 1952). At about the same time, work on the FTS was started at the Bureau of Mines in the US (Anderson, 1956, Anderson, 1984). In 1950, US firms built a commercial (365000 tons per year) ITS plant, developed by Hydrocarbon Research, in Texas (Keith, 1946). Syngas was obtained by reforming natural gas. The process used a fluidized bed reactor and, after operating troubles requiring the design of a new reactor, went on stream in 1953. Iron impregnated with about 1 wt% of K,CO, was used as a catalyst. The price of natural gas rose during this period and it became more profitable to sell the gas than to convert it to gasoline and chemicals; the plant, using what was called the Hydrocol process, was shut down soon after operation was deemed satisfactory. At about the same time, in South Africa, the Sasol 1 Fischer-Tropsch plant was built and commercial operation started in 1955. 4.5.2. On the chemistry of the Fischer-Tropsch synthesis (FTS) Schulz (1985) has provided a systematic definition of FIS chemistry that differentiates it from other reactions that involve the hydrogenation of CO: methanation, the synthesis of methanol, the Isosynthesis and the hydroformylation reaction. The Fischer-Tropsch synthesis (FTS) is catalyzed heterogeneously on metal catalysts which are extremely sensitive to poisoning by sulfur, on which CO is strongly chemisorbed, and which form metal carbonyls at high pressures but at temperatures too low for the FTS. These metal catalysts must be able to dissociate CO (split the carbon-oxygen bond). On certain catalysts under certain conditions, dissociative adsorption of CO is the main route; on Pd, Pt and Cu, associative adsorption is most likely, and on some catalysts (rhodium is probably the best example) CO is adsorbed in both ways. A better understanding of these observations can be gained from an examination of Fig. 17 (Brod&n et al., 1976). The dotted line at the left separates those metals (Cr, MO, etc.) that dissociatively chemisorb CO at room temperature. As the temperature is increased, the line is shifted to the right. At 200-300°C the borderline is between nickel and copper, rhodium and palladium, and osmium and iridium. It is fairly well established (Kellner and Bell, 1981; Katzer et al., 1981; Biloen and Sachtler, 1981) that metals to the left of the line catalyze

1. Wender/

Furl Processing Technology 48 (1996)

Mn

MO

Tc

Fe

,--_-: I i Ru

r----J w

1 Re

OS

i ROOM TEMPERATURE

229

I 1

I Cr

189-297

: Co

j cu r_____l

Rh ,---_: I j Ir

’ Pd a

Ag

Pt

Au

Ni

c SYNTHESIS TEMPERATURE 473-573 K

Fig. 17. Metals that adsorb CO dissociatively and non-dissociatively Tropsch reaction temperatures (BrodCn et al., 1976).

at ambient

temperature

and Fischer-

the ITS whereas those to the right are methanol synthesis catalysts. At still higher temperatures, however, the line is shifted further to the right, and at these temperatures CO is dissociatively chemisorbed even on copper and palladium. Poutsma et al. (1978) showed that Pd, Ir and Pt catalyzed the synthesis of methanol with high selectivity at higher synthesis gas pressures. The IT reaction yields a wide spectrum of hydrocarbons and oxygenated compounds. The major constituents of the hydrocarbons are paraffins and olefins, and primary alcohols are usually the chief oxygenated products. Straight-chain paraffins, along with some 2-methyl-branched paraffins, predominate among the saturated hydrocarbons; the major olefin components are terminal olefins. A large number of reactions occur during the IT reaction; the major ones are given below Paraffins(2n

+ l)H,

Olefins2nH,

+ nC0 -+ C,H,,

Alcohols2nHz

+ nC0 -+ CnH2n+2 + nH,O + nH,O

+ nC0 + C,H*,,+ ,OH + (n - l)H,O

WatergasshiftCO

+ H,O -+ CO, + H,

(4.23) (4.24) (4.25) (4.26)

Boudouard reaction 2C0 + C + CO,

(4.27)

Coke depositionH,

(4.28)

f CO + C + H,O

For catalysts that are good water gas shift (WGS) catalysts, such as alkalized iron, the water formed in Eqs. (4.23), (4.24) and (4.25) reacts with CO to form H,, so that the apparent HZ/CO usage becomes smaller. For a catalyst such as cobalt, which is not a good WGS catalyst, water is the main reaction product. The carbon number distribution of organic compounds is extremely wide and much effort has been expended in finding ways to improve the selectivity of the FfS for desirable products: gasoline, diesel fuel, C,-C, olefins and alcohols. Glycols are not produced in the FI reaction. Aromatic products become significant only at high reaction temperatures. When several reactions involving the same reactants are thermodynamically possible, the yields of various products depend upon their relative reaction rates and on the rates of ensuing reactions. At the temperatures usually employed in the ITS, the actual

230

I. Wender/Fuel

Processing Technology 48 (1996) 189-297

selectivity found differs very much from that expected from thermodynamic calculations. An excellent summary of the thermodynamics of the FTS has been given by Anderson ( 1984). Anderson (1984) and Frohning et al. (1982) have pointed out that data on the heats of reaction are important, as the FTS is strongly exothermic. Indeed, removal of the heat of reaction, about 25% of the calorific value of the syngas, is probably the chief problem in practical application of the synthesis. Excessive catalyst temperatures can lead to undesirable products, carbon deposition, catalyst deactivation or catalyst disintegration. The heats of reaction per carbon atom of the products vary only slightly with temperature, increasing with increasing carbon number for paraffins and decreasing with increase in the carbon number of the olefins formed. Enthalpy changes for reactions yielding hydrocarbons and CO, exceed those for the corresponding reactions yielding H,O. The enthalpy difference, about 9 kcal, is due to the WGS reaction, which is relatively independent of temperature. The heats of reaction for the formation of alcohols, also somewhat independent of temperature, in kcal mol- ’ per carbon atom are -23.9 for methanol and -29.5 for ethanol (Wagmann et al., 1982). The FTS is normally carried out at pressures from 0.1 to 4 MPa with c- 425°C as the upper temperature limit. As FT reactions involve a decrease in the number of moles, conversions at a given temperature increase rapidly with increase in pressure. The pressure and temperature limits of the reaction are determined by changes in selectivity and in the rate of catalyst deterioration. Anderson (1956) presented equilibrium data for the hydrogenation of CO to hydrocarbons and water, plotting AGO/n (AC per mole per carbon atom) versus temperature. The AGO/n curves for paraffins are about parallel but become more positive with increasing carbon number (Fig. 18). Below 573 K, AGO/n becomes more negative with increasing carbon number for olefins. In the FT temperature range, AC/n values for aromatic and saturated cyclic hydrocarbons, which are not significant FI products, are essentially the same as those for olefins of the same carbon number. Large equilibrium yields are possible for all primary, straight-chain alcohols except methanol. Methanol is formed in very small amounts in the FIS. As kinetic rather than equilibrium considerations are controlling in the FI’S, only limited amounts of useful information are furnished by thermodynamics. The products obtained depend largely on the selectivity of the catalyst. But certain conclusions derived from thermodynamics help us understand these syngas reactions Methane is the preferred product at all FIX conditions. * The relative stability of various compounds varies with the temperature of the FTS (25-425°C). The order is paraffins > olefins > alcohols. Although methane is the preferred paraffin, for olefins the favored products are compounds with the highest carbon number. Above = 425°C at 0.1 MPa this reverses and light olefins are favored. * The favored alcohol is the one with the highest carbon number. The equilibrium conversion of syngas increases with pressure. Owing to the effect of pressure (and temperature) on the activity and selectivity of the catalyst, the upper limits of usefulness for iron catalysts are 3-4 MPa and = 400°C. ??

??

??

231

-3

1

k

k

b

10 1

12 I

Carbon Number Fig.

18.Plot of AC/n

versus carbon number and temperature

(adapted from Stall et al., 1969; Janaf,

1971).

The water gas shift reaction is favored under FIS conditions; iron is a particularly good catalyst for this reaction. The Koelbel-Englehardt reaction (Koelbel and Englehardt, 195 11, the reaction of CO plus H,O to produce hydrocarbons, is thermodynamically more favored than the usual ITS from syngas. This reaction will be discussed later. The concentrations of n-paraffins and terminal olefins greatly exceed equilibrium values in the FTS. Subsequent isomerization reactions are not important. The reaction of ethylene and of ethanol with syngas is thermodynamically possible at FI temperatures. Incorporation of higher olefins or higher alcohols is less favored. Thermodynamically, methanol incorporation is favored over that of ethanol. The term “incorporation” denotes the building-in of an organic molecule together with CO andH, intheFIS. As with higher hydrocarbons and olefins, the amounts of oxygenated compounds (alcohols, aldehydes, acids, ketones) are also formed in much higher concentrations than predicted from thermodynamic calculations. However, these oxygenated compounds readily interact with each other under ITS conditions (Weitkamp and Frye, 1953). CH,CH,OH

e CH,CHO

(4.29)

+ H,

CH,CHO+H,O~CH,COOH+H,

(4.30)

CH,COCH,

(4.31)

Olefin hydrogenation

+ H, e CH,CHOHCH, and alcohol dehydration

are thermodynamically

favored under

I. Wet&r/

232

Fuel Processing

Technology

48 (1996) 189-297

FT conditions. Olefins and paraffins may be formed by dehydration of alcohols and hydrogenation of the olefins, respectively, and by primary reactions. . The Boudouard reaction, in which CO can form free carbon and CO,, is favored under all synthesis conditions. It is, however, possible to suppress this reaction in many FI reactions. It is difficult to arrange experiments amenable to analysis on a group of reactions of such a high degree of complexity as the FI reaction. Anderson (1956) has discussed the kinetics of the FIS, their goal being to find a fundamental rate equation relating the differential rate to the partial pressures or concentrations of reactants and of products. However, there are too many catalyst variables affecting the kinetics, including the method of preparation, composition and ageing, and also the operating conditions and transport phenomena involving the catalyst and the reactants, intermediates and products. Starch et al. (1951), Anderson (1956) Anderson (1984), Kuo (1984), Vannice (1982), and Dry (1981) have summarized kinetic expressions derived for nickel, cobalt, iron and ruthenium catalysts. Vannice (1975), using Al,O, as the support, reported that the activities of Group VIII metals declined in the order Ru, Fe, Ni, Co, Rh, Pd, Pt and Ir. With SiO, as the support, the activity declined in the order: Co, Fe, Ru, Ni, Rh, Ir and Pd (Vannice, 1977). Most studies, however, have been carried out using small laboratory units, usually at low levels of conversion. Dry (1981) Dry (1996) reinvestigated the FTS over iron catalysts in large pilot plant reactors over wide ranges of conditions at high levels of conversion (including commercial conditions). He found that, at low levels of conversion, the rate was directly proportional only to the partial pressure of H,; at higher levels of conversion, water vapor pressure had a strong negative effect, but CO, had only a small influence on the rate equation r=mPcoPH2/(Pco+~PH*o) This resembles the equation proposed by Anderson (1956) and fits Sasol research findings. When PHZO is low (at low conversion levels), the equation becomes r=mPH2

(4.33)

With the full rate equation plus the WGS, the reaction profiles of pilot and commercial-scale reactors (fixed and fluidized bed units) could be accurately simulated in a computer model (Dry, 1981). Satterfield (1991) has proposed a somewhat different version of the FI rate equation. 4.5.3. On the mechanism of the Fischer-Tropsch synthesis (FTS) The intrinsic kinetic feature of the FTS is stepwise chain growth, in effect a polymerization of -CH,- groups on the catalyst surface. This is valid regardless of the products that are formed, paraffins, olefins or alcohols. There has been much research on chemisorbed intermediates in the FTS, reflecting on progress in experimental methods, work that has added to our understanding of the synthesis. Although it may appear that the FI proceeds by simple polymerization of methylene groups, a large number of identified or assumed species exists on the catalyst surface during reaction, some of which are depicted in Fig. 19 (Schulz, 1985).

1. Weruler/Fuel

H,

Processing

Technology 48 (19961 189-297

H, ,,OH, ,OH

,CHTR

R.

,O

Rt

0

c

c

0

C

CH-0

u

L8

19

z!

21

22

Fig. 19. Some surface species of Fischer-Tropsch

233

CO hydrogenation

(Schulz,

1985).

The large number of these surface species has led to the consideration of several mechanistic pathways for the FTS (Fischer and Tropsch, 1926; Starch et al., 1951; Anderson, 1956, Anderson, 1984; Pichler and Schulz, 1970; Dry, 1981; Biloen and Sachtler, 1981). However, there is general acceptance that chain growth in the FTS proceeds by a stepwise process. Herington (1946) first introduced into FT studies the probabilities of chain growth and chain termination, terms common to polymer chemistry. He considered the paraffins and olefins formed on a cobalt catalyst and postulated that they were formed by stepwise addition of a methylene (-CH,-) entity to the growing chain on the surface of the catalyst. Anderson (1956), Anderson (1984) and Friedel and Anderson (1950) analyzed the product distribution of a large number of FT runs using different fixed bed catalysts. They found that plots of 1ogWJn against the carbon number n yielded straight lines over a fairly large range of products (W,, is the mass fraction of a particular product). This showed that the probability of chain growth (Y was essentially constant. Most FT mechanisms assume that the monomer unit is the same weight wherever it is found in the chain. From this work an equation for mass fraction can be written as W,=n(l

(4.34)

-*)2&

This expression is equivalent to the Schulz-Flory equation (Flory, 1950; Schulz, 1935) which treats the Ff product distribution as a polymerization process. Eq. (4.34) is now generally referred to as the Anderson-Schulz-Flory (ASF) equation. It is usually written in the logarithmic form log-

w, n

(1 - cr)2 = n log (Y+

(4.35) a

234

I. Wender/

Fuel Processing Technology 48 (1996) 189-297

It is evident that, if Eq. (4.35) for the FI product distribution holds, the calculated value of cy from the slope (log a) should be consistent with that calculated from the intercept, log(1 - cu>*/(u. The probability of chain growth cx is the ratio of the chain propagation rate constant to the chain propagation plus the termination rate constants. The ASF equation has the consequence that methane can be synthesized in 100% selectivity; all other products have well-defined maxima in allowed selectivities as shown in Fig. 20 (Mills, 1993). Numerous FI studies leave no doubt that the ASF equation predicts the proportions of methane, to gasoline, to diesel, and waxes that will be produced in FT reactions. The highest selectivities attainable by the ITS are, in wt%, methane 100; ethylene 30; C,-C, olefins 56; gasoline 48. The large array of products from the FTS, as in the Sasol operation, includes a wide range of hydrocarbons mixed with some oxygenated compounds, necessitating a considerable number of separation steps. In wax production in fixed bed reactors, there is a discernible change in the slope of (Y plots around C ,,,, giving two (Y values, with the second always higher than the first. The result is an increase in the amount of heavy ends produced. In slurry phase operations, there is a large reservoir of liquid phase and the two (Y values are clearly observed in this case (Dry, 1996). There has been an enormous amount of research on the mechanism of the FI reaction. The earliest postulation was made by Fischer and Tropsch (1926), who suggested that carbon deposited from CO as a surface or bulk carbide eventually formed FT products. Pichler and Schulz (1970) postulated a mechanism whereby CO was inserted into a metal-alkyl or metal-hydrogen bond. Biloen et al., 1979 presented evidence for the stepwise insertion of CH, units produced from a syngas mixture. Ponec (1984) and Rofer-DePoorter (1981) have pointed out that the FT mechanism includes various pathways with the same intermediate leading to different products,

01

oa

0.3

Fig. 20. Plots of calculated selectivities (percent function of the probability of chain growth.

0.4

carbon

05

0.6

0.7

as

03

atom basis) of carbon

Lo

number

product

cuts as a

I. Wender / Fuel Processing

Technology 48 (1996) 189-297

235

whereas different intermediates can give the same products. This is probably the best view, although the work of Biloen et al. mentioned above and that of Brady and Pettit (19801, Brady and Pettit (1981) support the view that the principal FT mechanism involves a -CH,stepwise polymerization. Many mechanisms have been postulated for the FTS; a plausible one is given in Fig. 21 (Dry, 1996). The first step is the dissociation of CO, perhaps assisted by the chemisorption of hydrogen. Aldehydes and alcohols result from the insertion of CO into the chain; the latter is a chain growth termination event. Fairly simple catalyst and process changes in the FTS can yield large increases in oxygenated compounds of all chain lengths. In line with the chain termination steps, alcohol selectivity with iron catalysts is directly proportional to the hydrogen partial pressure, as is the olefin to paraffin ratio. There is a good correlation between the partial pressure of CO, and the selectivity to acids. Graphitic carbon forms by agglomeration of carbon atoms that are formed by CO dissociation. The effect of temperature on selectivity is consistent for all FT catalysts. As the temperature of the synthesis is increased, the methane selectivity increases, the amount of olefins in the product rises and the selectivities toward oxygenated compounds decrease. The temperatures used at Sasol bear out these effects and take them into account in actual operation. Promoters have been divided into two types according to their mode of action. Oxides that are difficult to reduce, such as SiO,, Al,O,, MgO, ThO,, La,O, and ZnO, are called structural promoters. They furnish a large surface area and prevent recrystallization and sintering of the active catalyst. There is evidence that these so-called textural promoters often interact chemically with various oxidation states of the catalyst and can exchange oxygen atoms. Chemical promoters exert their influence by mechanisms not clearly understood. They may transfer electrons to the catalysts or even block pores. Alkalis and their salts are the chemical promoters most often used-they are basic salts and catalyst basicity, especially for iron catalysts, is a key parameter. Promoter effects depend not only on the type and amount of alkali salt added but also on the interaction of the alkali with the support, with other promoters, and with impurities. Potassium oxide, K,O, is the chemical promoter used most often. If the alkali reacts with the support, the basicity will be decreased and more alkali must be added. Alkalis are considered to be electron donors. The effects of alkali promotion on FT catalysts such as iron may be summarized as follows: (a) suppression of hydrogenation capability, (b) increase in CO dissociation, (c) increase in formation of long-chain hydrocarbons, and (d) decrease in the conversion activity of CO. The sample ratto P,,/Pco adequately represents the selectivity for a fixed bed reactor operating at a low temperature with an extruded iron catalyst (Dry, 1990). The partial pressure of CO, and also the total pressure (Dry, 1981) influence the selectivity for iron catalysts which operate at a higher temperature. IT catalysts can lose activity for a number of reasons: (a) sintering (loss of active surface area due to growth of crystals), (b) conversion of the active parts of the catalyst

236

1. Wender/Fuel

Processing TNlTlATION

Technology 48 (1996) 189-297 AND C COMPOUNDS I

GRAPHITE t +c

M

M

M

I

1

M

H 2

H2

1

CH2

LcHq

5 M M

CHAINGROWTH

(1)

C9

-

CH2

f

CH

.-__C)

-

i M

5

M

M R /

(2)

k CH I2 Cl-l (A)

7H3

CH

2 H2

-=W W

f M

1:

ZHZ,

RCH2CH2COOH

I

.c.

RCH2CH0

Fig. 2 I. Mechanism of the IT reaction (Dry, 1990).

to inert phases (i.e. a metal to a metal oxide), (c) deposition of carbonaceous substances on the active surface area (carbon deposition), or (d) chemical processing or poisoning of the surface (sulfur is the chief culprit here). Poisoning by sulfur compounds apparently occurs chiefly by adsorption of sulfur atoms on active metal sites, destroying their catalytic activity. The amount of allowable

I. Wender/

Fuel Processing

Technology 48 11996) 189-297

2.17

sulfur in the ET reaction is on the order of a few parts per billion. Coke deposition is probably the most important mode of catalyst deactivation. Tungsten and molybdenum catalysts that are more sulfur resistant than the usual FT catalysts have been found (Murchison and Murdick, 198 1; Bartholomew, 1991). In general, they have low activity and poor selectivity. An alkalized molybdenum catalyst that is selective to low molecular weight products has been developed. Bartholomew (1991) has provided an excellent review of recent developments in catalysis of the FT reaction. 4.5.4. Circumventing Fischer-Tropsch chain growth kinetics. Use of zeolites A great deal of work has been done on circumventing the ASF “constraints” but these effects have usually been shown to be short-lived or the result of errors in analysis or data interpretation. But interesting and somewhat effective methods have been examined as a way of circumventing ASF chain growth kinetics. They involve interception of FT intermediates by two general approaches: (a) a FT catalyst either supported on a zeolite or physically admixed with a zeolite or (b) a multi-step (generally two-step) process involving FlS followed by an upgrading step using a zeolite catalyst. Although many zeolites have been used, most work has centered on Mobil’s ZSM-5 shape selective catalyst. FT catalysts usually operate in the 250-350°C temperature range, whereas zeolites such as ZSM-5 are generally employed at somewhat higher temperatures, so that optimum conditions for FT catalysts and zeolites differ. Reactions with combined catalysts must be carried out at intermediate temperatures. Haag and Huang (19791, Haag and Huang (1981) demonstrated the advantages of using iron-potassium or cobalt in the first stage of a dual reactor arrangement with ZSM-5 in the second conversion stage. In this way, the ZSM-5 catalyst can be operated at about 355°C and the FT catalyst at normal FT temperatures. It is also simpler to regenerate each catalyst, as each catalyst requires a different regeneration procedure. Nijs et al. (1979) were the first to use zeolites as a chain-limiting catalyst in the FE. The FTS was conducted with the usual Ru/SiO, catalyst and gave an ASF distribution with 60% of the product above C ,2. A combination of a FT catalyst and a faujasite type, RuNa zeolite gave a product with less than 1% above C ,2 but with a marked increase in unwanted selectivity to methane. Ballivet-Tkatchenko and Tkatchenko ( 198 1) prepared catalysts by thermally decomposing metal carbonyls of Fe, Co and Ru in the cavities of zeolites; they obtained selective formation of C , -C, hydrocarbons and the usual ASF plot. Nijs et al. (19791, Rao and Gormley (1990), Bartholomew (1991) and others have discussed possible explanations for the above effects. Shape selectivity seems obvious but metal dispersion effects (Shamsi et al., 1986) and experimental occurrences such as liquid product holdup in the support material may play a part. It is possible that selective adsorption of heavier molecules takes place on the zeolite and apparent deviation from ASF kinetics only occurs in about the first 24 h of reaction. Mobil has developed a two-stage slurry FT/ZSM-5 process combining slurry-phase FT technology with a fixed bed ZSM-5 reaction (Kuo, 1983). Based on a pilot plant study, a conceptual design for a 27000 bpd gasoline plant was drawn up. The product

238

I. Wet&r/

differs from the usual FI 1985). 4.6. Fischer-Tropsch

Fuel Processing

product

Technology

in having

48 (1996)

significant

189-297

amounts

of aromatics

(Kuo,

Sasol plants

As mentioned earlier, the first commercial FT plant built for profit was the iron catalyzed, medium pressure (loo- 150 psig) fixed fluidized bed reactor in the Hydrocol Process at Brownsville, Texas. It used syngas made from natural gas and operated from 1950 to 1953. At about this time, 1955, in South Africa, the Sasol 1 FI plant, based on the use of iron catalysts, went into commercial operation. This country has an abundance of coal, a dearth of petroleum resources, and large centers of population far from the ocean, together with a political and economic need for resource autonomy. The Sasol acronym was derived from the South African Coal, Oil and Gas Corporation, Ltd. a private company established with government funding through the Industrial Development Corporation to convert coal to liquid fuels and chemicals via the ITS. Sasol 1 produced about 8000 bpd of products, supplying about 5% of South Africa’s motor fuel needs plus other fuels and chemicals. Sasol 2, with an output of 50000 bpd, started operation in 1980; Sasol 3 went on stream in 1983. South Africa then had the capability of providing = 40% of their liquid fuel and petrochemical needs from coal using the FT process. An excellent account of the PI’S has recently been given by Dry (1996). Sasol 1 initially had two types of reactors. The fixed bed tubular reactor, termed the ARGE reactor, used a precipitated iron catalyst promoted with copper and a potassium salt such as K2C0,. The other type was an entrained fluidized bed reactor, the Synthol reactor, which used a fused iron catalyst with alkali and other proprietary promoters. Sasol 2 and 3, however, use only Synthol reactors each consuming about 40000 tons of coal per day. Sasol 2 is the largest grass roots, single-purpose processing complex in the world, occupying an area of one mile by one and one-half miles and valued at several billion US dollars. Sasol 3 is essentially a twin plant located adjacent to Sasol 2. The Sasol plants are specific applications of the ITS. Coal is gasified in Lurgi gasifiers which yield significant amounts of methane and large amounts of CO, in the gaseous products and HJCO ratios that hover around 2. Lurgi gasifiers are generally best used with lower rank, non-caking coals. The overall methane yield from the gasifier and the FTS itself can reach as high as 20%. Early on, no pipelines for the distribution of the methane existed in South Africa and the amount of methane made exceeded the country’s demand; the Sasol plants were forced to reform the excess methane to make more syngas, which was then recycled. As a result, the thermal efficiency at Sasol, defined as the lower heating value (LHV) of the products divided by the LHV of all the coal used, was only about 40% for the production of motor fuels, including gasoline, diesel oil and LPG (Eisenlohr and Gaensslen, 1981). The efficiency could rise to about 60% if the methane could be utilized as such. Research on the FTS decreased abruptly in the 1950s. Only South Africa had a modest program devoted mostly to practical problems that arose from operation of Sasol 1, while a small group continued research at the US Bureau of Mines (Anderson, 1956, Anderson, 1984).

1. Wrnder / Fuel Processing

Technology

48 (19961 189-297

239

The Arab oil embargo of 1973 furnished great impetus for new research on the Ff’S with emphasis on improving gasoline and diesel oil yields. Sasol workers had greatly improved their Lurgi gasifier operation, effectively cleaned the gases that emerged from the gasifier, removing CO, and almost all the sulfur compounds, and developed excellent methods for the separation of the multitude of products obtained. They were able to maximize gasoline and diesel production while separating out many “petroleum” feedstocks such as ethylene and propylene and also oxygenated chemicals (ethanol, acetic acid, acetone, etc.>. The then projected distribution of products from Sasol 2, in tons per year, was: motor fuels 1500000; ethylene 185000; chemicals 8.5 000; tar products 185 000; ammonia (as N) 100000; and sulfur 90000. Some 2 140000 tons per year of saleable products were envisioned. The plants have over the more than 15 years of operation undergone a large number of additions and expansions and trhe Secunda plants now produce nearly 7 million tons per year of marketable fuels and chemicals (Geertsema, 1996). Under certain conditions, synthetic fuels from coal are viable, and improvements, some quite striking and valuable, are continuously being made in the FTS. Much was learned from the building of Sasol 2, so that the cost of Sasol 3 was significantly less than that of Sasol 2. The Sasol Fischer-Tropsch process for the synthesis of fuels and chemical is outlined in Fig. 22 (Dry, 1996). More than 100 products are marketed by Sasol today, furnishing a wide spectrum of fuels and chemicals. A complex refinery system is required to deal with all the products. Sasol, originally funded with Government money, was fully privatized in 1979 and its shares have since been traded on the Johannesburg Stock Exchange. The only protection that was applied was in regard to gasoline sales; this came to 4.3 cents per gallon to the end user. It is estimated that this tariff protection saved South Africa $1.6 billion a year in foreign exchange; it furnished Sasol the difference between $21.40 a barrel of oil and the world price of $18.77 at the end of 1995. However, the government announced that it would reduce the tariff protection in 1996 to provide $19 a barrel and cut it further by mid-1999 to assure $16 a barrel. However, there is no protection for chemicals, which must compete in the international market. When the Sasol 2 and 3 plants began operation, synfuels provided about half of South Africa’s needs. With growth in other areas in the last decade. Sasol now supplies only a third of the country’s fuels. It is informative to look at Sasol’s operating profit contributions (Table 6). The substantial increase in profits from petrochemicals is due largely to capital projects specifically aimed at increasing the number and variety of output of high-value chemicals. Sasol operations have a significant effect on the economy of South Africa. More than a trillion (10’2> dollars is saved on its foreign exchange and a very large number of jobs are created, directly or indirectly. 4.6. I. Sasol reactors Sasol mines a total of about 40 million tons of coal per year. There are three major plants, Sasol 1 at Sasolburg and Sasol 2 and 3 at Secunda. The total of 97 Lurgi gasifiers

240

I. Wender/

Fig. 22. Block diagram from Dry, 1990).

for Sasol plant processes.

Table 6 Source of operating

Fuel Processing

profit at Sasol (1994-95

Technology 48 (1996) 189-297

(Reprinted

with permission

of Elsevier Science Publishers

profit R 2805 million = $780 million) Percentages

Synfuels Coal Refining, fuel gas Petrochemicals

93-94

94-95

48 14 21 17

44 11 14 31

loo

loo

I. Wender /Fuel

Proce.s.sing Technology

48 (1996) 189-297

241

consumes 27 million tons of coal per year. Gasifier coproducts include ammonia, sulfur, phenols, cresols, pitches, anode grade coke and, soon, metallurgical grade coke. Two types of reactors were used at Sasol 1 until recently. They were the ARGE fixed bed reactors. which are still in use, and the Synthol circulating fluidized bed (CFB) reactors, which were shut down at Sasol 1 in 1992 but are still in full operation at Sasol 2 and 3. The fixed bed catalyst uses fine metallic iron filings as the raw material; it is a precipitated and extruded catalyst with alkali promoters. Catalyst changes are made after 70-100 days. Low temperatures, about 225°C at 25 atm (one at 45 atm since 1984) are used. At higher temperatures, carbon deposits on the catalyst would lead to reactor plugging. The heat of reaction is removed by circulating water on the outside of the tubes. The Synthol catalyst is made from millscale from a nearby steelworks. The ground catalyst with added promoters is then fused in an open arc furnace. The catalyst leaves the furnace at about 15OO”C, solidifying as it cools. The ground catalyst ( = 200 mesh) is reduced with hydrogen and stored in the absence of air. In the Synthol reactors, which operate at about 340°C and 25 atm, flowing hot catalyst from the standpipe is entrained by the feed gas into the reactor zone where the heat of reaction is transferred to heating coils. The turbulent flow of gas and catalyst passes through heat exchangers to the wide settling hopper above the standpipe; here catalyst and gas disengage. Hydrocarbon vapors and gas leave the reactor via cyclones to remove entrained catalyst, returning them to the settling hopper. About a third of the generated heat is removed by internal heat exchangers; the rest leaves with the recycle gas. Temperatures of the exit gas are = 320-360°C. As depicted in Table 7 (Jager et al., 19901, the Synthol reactor produces more light hydrocarbons, more olefins, more oxygenated compounds, more gasoline and less heavy oil and waxes than the fixed bed. Sixteen later versions of the CFB reactors were erected at Sasol 2 and 3 in 1980 and 1983 at Secunda. They had capacities some three times that of the corresponding reactors at Sasol 1, to an equivalent of 6500 bpd. The CFB reactor, however, has a number of limitations (Jager et al., 1990) - It is physically complex and is suspended in a complex structure. - Circulation of large tonnages of catalyst results in considerable recycle gas compression with added costs. Table 7 Product selectivities Jager et al., 1990)

of Sasol commercial

Fixed bed

Product

CH, C, -C, C, -C,

reactors. (Reprinted

oletins paraffins

Gasoline Middle Distillate Heavy Oils and Waxes Water Soluble Oxygenates

with permission

of Baker

A.G. Publishers

Synthol (fluidized bed)

4

7

4

24

4 I8 I9 48 3

6 36 I2 9 6

from

242

I. Wender / Furl Processing

Technology 48 (1996) 189-297

Carbon formation increases with increase in temperature and lowering of the HZ/CO ratio. Carbon deposition reduces the amount of iron catalyst in the reactor, limiting reaction rates and catalyst life. * There is considerable erosion at the various bends in the reactor. + No further scale-up of the Secunda CFB reactors is deemed feasible. A development at Sasol led to the successful commissioning of a commercial scale conventional fixed fluidized reactor as an alternative to the CFB reactor. This reactor type is now called the SAS (Sasol Advanced Synthol) reactor. A 5 m diameter reactor operated in Sasol 1 from 1989 to 1992. During this time, important design data were gathered. With the revamping of the Sasol 1 facility, it was decided to dedicate it to the production of waxes and waxy products. Therefore, this reactor was transformed into a new type of reactor for the production of heavy hydrocarbons. A new SAS reactor has been recently commissioned in Sasol 2, Secunda. This larger reactor, 8 m in diameter and 38 m high, is coupled to the existing production facility. The operation during the initial 6 months has been smooth. In the SAS (Sasol Advanced Synthol) reactor, the gas enters the reactor via a distributor and bubbles through the catalyst bed. It is referred to as “fixed” since the bed, although fluidized, is not transported as in the CFB reactor. The main benefits of the SAS reactor are * Cost, less than half that of the older reactors. - Better thermal efficiency. - Pressure drop over the reactor is lower so that gas compression costs are lower. - Isothermal behavior. * Greater flexibility. - Significant savings in operating and maintenance costs. - High oil selectivities are achieved and the percentage conversion is higher than in the CFB reactor. - Scale-up to about a 16500 bpd SAS unit is envisaged. A preliminary economic analysis and comparison of the SAS and CFB reactors indicated that the SAS will be a leading contender for future synthetic fuels from coal-generated syngas. The ability of the FTS to co-produce chemicals and to produce fuels with low sulfur and aromatics makes this an attractive option. SAS reactors are less than half the size of the CFB reactors of the same capacity (Jager et al., 1990). In May 1993, Sasol brought into operation its fourth and newest type of reactor, a first-of-its-kind commercial large-scale slurry bed reactor know as the SSBP (Sasol Slurry Bed Process; Fourie, 1992; Geertsema, 1993). It resembles a SAS reactor except that the catalyst is suspended in a liquid, usually a FI wax. The 5 m diameter, 22 m high SSBP has been operating successfully since start-up. Although great potential exists to vary product yields and selectivities by adjusting operating conditions or catalyst properties in the SSBP, its first use is in the production of FT wax. The production of high-quality diesel fuel is a longer term goal, although the wax can be cracked to yield this type of liquid fuel. The SSBP has a number of advantages including a low pressure drop, isothermal behavior, good scale-up potential, on-line catalyst removal, improved catalyst economy and low turndown ratio. Construction of a slurry-phase reactor is about 45% cheaper than that of the same capacity multi-tubular unit (Jager et al., 1994). ??

I. Wen&r/

Fuel Processing

Technology 48 (1996) IN-297

243

ling

Boiler Feed Water

Synthesis Gas

SYNTHOL REACTOR

Fig. 23. High temperature

SASOL ADVANCED FLUIDIZED REACTOR

IT reactors.

Of the four Sasol commercial reactor options, two are operated at high temperature (= 34O”C), Fig. 23, and two at low temperatures (220-27O”C), Fig. 24 and Table 8. The high temperature process (HTFI) was first used at Sasol 1, 2 and 3 for circulating fluidized reactors (Synthol units). An updated variant, a more classical fluidized bed with smaller costs and larger benefits (the Sasol Advanced Synthol process), was operated at Sasol 1 and now at Sasol 2 Secunda. It is likely that any new commercial venture will use either the SSBP or the SAS reactor, depending on the type of product that is desired. The high temperature process yields large amounts of olefins, a lower boiling range and very good gasoline. Diesel fuel can be produced readily by oligomerization of olefins. Substantial amounts of oxygenates are also produced. The low temperature FI (LTFI’) has, until recently, been commercially operated only in fixed bed tubular reactors using an iron catalyst or, as Shell now operates, with a modified cobalt catalyst. This process yields much more paraffins and linear products and can be adjusted to very high wax selectivities. The primary diesel cut and wax cracking products can give excellent diesel fuels. The very linear primary gasoline fraction needs further treatment to attain a good octane number. Olefin and oxygenate levels are lower than for high temperature FT reactors. As mentioned, Sasol has commercialized a low temperature FI process using the concept of a slurry bed. Iron based catalysts are used in the SSBR. The single slurry bed unit at Sasol 1 now uses all the syngas previously fed to three decommissioned Synthol units. Synfuels production has ceased at Sasol 1. The major products from Sasol 1 are waxes from the new slurry unit and from six tubular ARGE reactors. A range of

244

I. Weruler / Fuel Processing

Technology 48 (1996) 189-297

Fresh

Tube Bundles

-

Steam TO

-

Gas

Outlet Synihesis -km

ARGE

REACTOR

SLURRY

Fig. 24. Low temperature

BED REACTOR

(SSBP)

FT reactors.

products including very high purity ethanol and also propanol, paraffin cuts, phenol, o-cresol, ammonia and industrial gas is also produced. Typical product distributions from three reactor types are given in Table 9. Of the 40 million tons of coal per year used at Sasol, 65% goes to gasification and 35% to power/steam. The major products (mass %) are gasoline 38, diesel 22, industrial gas and LPG 13, ethylene and propylene 4.4, polypropylene 2.0, alcohols and ketones 6.3 and “black” products 5.8 (Geertsema, 1993). Iron catalysts have been used in all Sasol reactors. However, variously promoted forms of cobalt have been tested in pilot plants. Potentially, a cobalt catalyst should work well in the SSBR reactor. Cobalt generally gives more saturated products (less olefins) and also fewer oxygenated compounds.

Table 8 Sasol Fischer-Tropsch commercial reactors. CFB = circulating fluidized bed; SAS = Sasol advanced ARGE = low temperature fixed bed reactor; SSBP = Sasol slurry bed reactor Geertsema, 1993

Current capacity Capacity per reactor Expected scaled-up capacity per reactor Max % gasoline Max % diesel

synthol;

CFB HTFT

SAS HTFI.

ARGE Lm

SSBP LTFI

110000 6500-7500

3400 operated, 11300design 16500 80 70

3200 500-700

2400 2400

1200 (design) 30 85

10000 30 85(?)

80 70

I. Wender/

Table 9 Typical primary Fischer-Tropsch

Fuel Processing

Technology

product spectra (Geertsema, ARGE

Paraffins Oletins Aromatics Oxygenates % n-Paraffins

in paraffin cut

48 (1996)

189-297

245

1993)

SSBR

SYNTHOL

= SAS

cs-Cl2

cl,-Cl,

cs-Cl2

c,,-c,s

c5-c,0

c,,-c,4

53 40 0 7 100 95

65 28 0 7 100 93

29 64 0 7 100 96

44 50 0 6 100 95

13 70 5 12 100 55

15 60 I5 IO 100 60

4.6.2. Petrochemical expansion at Sasol Sasol is investing about $1 billion in plants commissioned in 1993 and planned to 1995 and beyond. Foreign exchange savings should reach about $1.6 billion per year. Some of Sasol’s commitments are listed in Table 10. Product streams from Sasol plants contain a wide variety of terminal linear olefins, with C,-C, exceeding 750000 tons per year. Of this, more than 250000 tons per year consists of 1-pentene and more than 170 000 tons per year is I-hexene. Currently these olefins are utilized in the gasoline pool. Sasol initiated the first phase of a large olefins project (Chemical Marketing Reporter, 1992) with a goal of a capacity of 450000 tons per year of I pentene and 1-hexene. The Sasol plants employ a newly developed, low cost recovery and purification process which will produce copolymer grade I-hexene and I-pentene at a combined capacity of 100000 tons per year. The first phase was commissioned in 1994. 1-Pentene copolymers are expected to provide interesting advantages in improving the performance of I-butene copolymers while meeting food grade restrictions on higher olefin copolymers. It also could be an excellent comonomer for ultra low density polyethylenes and elastomeric polyethylenes. In addition to the utilization of these olefins in polymer production or perhaps in their conversion to alcohols via the hydroformylation (0x0) reaction, Sasol has a number of

Table IO Current Sasol expansion

(1994)

Proiect

Commissioned

$ million

Sasol I upgrade Acrylic fibres Anode coke Open cast mine, Sasol I Cresylic acids a-Olefins Increased industrial gas + 0, New gas pipeline Alkylamines Acrylonitrile SAS reactor Secunda

5/ I993 6/1993 6/ I993 4/ I993 4/ 1993 3/1994 7/ I994 7/ I994 9/ I994 l/1995 9/ I995

215 120 100 40 27 122 100 47 I6 102 41 990

I. Wender/Fuel

246

Processing

Technology48

(1996) 189-297

further products under consideration. These include acetic acid and acetates recovery, methanol synthesis, MTBE and ETBE, bisphenol A, natural gas utilization as a source of syngas, and exploration for oil. 4.63. Natural gas to FT products Natural gas is well known as a clean, efficient hydrocarbon source, World proven natural gas resources amount to some 4000 trillion (4000 X lO’2) cubic feet, but about half is found in areas far from markets and it is not economic to transport it as a gas for long distances. The world is dependent on liquids for transportation fuels and the conversion of remote natural gas to clean liquids is a highly desirable goal. As an example of the possible use of this remote resource, gas conversion technology could furnish about 100000 bpd of premium transportation fuels for some 20 years from only six trillion cubic feet of gas reserves (Velocci, 1991). Some driving forces for the utilization of natural gas, especially of remote gas, are (a) long pipelines are not only costly but often simply not feasible, (b) liquified natural gas (LNG) markets are limited, (c) there is promise in utilization via the synthesis of methanol or a ET process, and (d) political and environmental issues are driving the development of processes for conversion of natural gas to fuels and chemicals. There is considerable interest in cobalt IT catalysts and a large number of patents have been issued to major oil companies, including Gulf, Exxon, Shell and Statoil. The cobalt catalysts under development by the various companies can be grouped as follows (Goodwin, 1991); they generally contain about 20 wt% of cobalt (Table 11). These companies were or are focusing on the production of heavy ends via the ITS; the products can then be hydrocracked to yield desirable middle distillate fractions. Although Fe, Co, Ni and Ru are the most active metals for the FTS (Vannice, 1982), only Fe and Co are feasible catalysts. When the source of syngas is coal gasification, the high CO/H, ratio so produced favors the use of an iron based catalyst. Iron has high WGS activity so that less hydrogen is needed. However, there is one CO, molecule produced for each -CH,in the products when iron is used and this may be a consideration in regard to the so-called greenhouse effect. On the other hand, cobalt is not a good WGS catalyst and oxygen exits mostly as water. Addition of small amounts of a noble metal can lower the reduction temperature needed for IT cobalt catalysts (Raab et al., 1990). Fiato et al., 1989 have suggested that the function of the second metal is to improve the regenerability of the catalyst. The noble metal probably acts as a source of spillover hydrogen at lower temperatures,

Table 11 Typical catalyst constituents Company

Primary

Secondary

Gulf Exxon Shell Statoil

co co co co

Ru Re/Ru with or without noble metal Re

1

Secondary oxides oxides zfl2

oxides

2

support alumina titania silica alumina

1. Wender / Fuel Processing

Technology 48 (I 996) 189-297

247

resulting in better reducibility or removal of carbonaceous material from the catalyst by hydrogenation. Important catalyst parameters given in the patents for improving catalytic properties include the method of preparation, impurities in the support, support porosities, Fe or Co dispersion, metal-support interaction and phase transformations. 4.6.4. The Mossgas gas to fuels process In 1987, a project to produce liquid fuels from offshore natural gas in South Africa was announced. The government-supported project, known as the Mossgas project, had the Industrial Development Corporation as a major shareholder. The project is located in Mossel Bay in the Eastern Cape area of South Africa (van Rensberg, 1990). The natural gas is 72 miles offshore from Mossel Bay. An underwater pipeline was built to transport about 168 million cubic feet of gas per day and 50 metric tons per hour of gas condensate to the town of Mossel Bay, the offshore gas to be recovered with separation of the associated condensate. The total was then to be converted to transportation fuels. Sasol licensed the use of three Synthol (CFB) reactors to the Mossgas project. This first use of the Sasol Synthol process on a natural gas feed was commissioned in 1993 but the current project is now catalogued as a “commercial mistake.” The gas field would be depleted by 1997 and the $3.3 billion that it cost would not be recuperated. The South African government has given the go-ahead for Mossgas to invest a further R 360 m ($100 million) in satellite gas fields. There is the possibility to convert the facility into a fully fledged petrochemical producer. It is hoped that a buyer for the plant would be found (Cohen and Ensor, 1996). 4.6.5. The Shell middle distillate process The Royal Dutch/Shell companies have been carrying out research and development since the late 1940s on converting coal and natural gas to liquid fuels. Impetus to their work was furnished by the 1973 oil crisis. They developed the Shell Coal Gasification process (SGP) for power generation via syngas; a 250 MW unit, operated by the Dutch electricity generation board, began operation in 1993. In 1989, Shell, together with its partners Petronas, the Sarawak State Government and the Mitsubishi Corporation, announced the construction of a commercial Shell Middle Distillate Synthesis (SMDS) plant in Bintulu, Sarawak, Malaysia to convert natural gas to liquid transportation fuels (van der Burgt et al., 1990; Sie et al., 1991; Eilers et al., 1990; Ansorge and Hoek, 1992; Tijm et al., 1993). The aim is to convert 100 million cubic feet of natural gas from offshore fields to about 500000 tons per year of hydrocarbons. Natural gas reserves are large and growing, rivaling the level of crude oil reserves. Capital expenditure for a syngas plant based on coal is about twice that of a plant based on natural gas. The SMDS process involves three stages: syngas manufacture, heavy paraffin synthesis (HPS) via the FTS, and heavy paraffin conversion (HPC). The products, mainly kerosene, gas oil and some naphtha, are finally separated by distillation. The high quality of the products, with no sulfur or aromatics (typical of FT products), coupled with an anticipated growing demand for middle distillate fractions, especially in the developing countries, would seem to make the SMDS a desirable route to environ-

248

I. Wender / Fuel Processing

Technology 48 (1996) 189-297

mentally acceptable liquid transportation fuels. The SGP gasification to syngas uses a non-catalytic autothermal partial oxidation of methane operating at 1300-1500°C at up to 70 atm with a carbon efficiency of over 95%. Little adjustment of the desired 2:1 HJCO ratio is required. CH, + l/20, --) 2H, + CO + - (CH,) - -t H,O (4.36) Steam reforming of the C ,-C, FI products yields additional hydrogen. In the next step, the heavy paraffin synthesis (HPS), the reaction mechanism follows ASF polymerization kinetics, characterized by the probability of chain growth (Y versus chain termination. A high (Y(Fig. 25) corresponds to a high average molecular weight, highly linear paraffinic product. Shell has a proprietary catalyst based on cobalt, probably plus a noble metal such as ruthenium or rhenium. The desired product is a long chain hydrocarbon wax. The reaction is carried out in a fixed bed tubular reactor with efficient energy recovery. The catalyst is expected to have a useful life of over a year and is regenerable. The FT catalyst and operating conditions were selected so as to yield a heavy product with a high cx value; the formation of light hydrocarbons is minimized. The HPG product is fractionated and the fraction boiling above the gas oil range is recycled to the HPC reactor. Selectivities are influenced by varying process severity or conversion per pass. Products can be about 50% kerosene on total liquid product or, for the gas oil mode of operation, some 60% gas oil. The combination of chain length independent FTS of heavy paraffins and their selective cracking is the basis of the SMDS process. Both the kerosene and the gas oil

0.85 a =

0.90

PROBABILITY OF CHAIN GROVi-lH

co (cLAssICAL) Fe (CLASSICAL)

_------a-------_---_

Fig. 25. Molecular mass distribution

in raw product (Tijm et al., 1993).

0.95

I. Wender/

Fuel Processing

Technology 48 (19961189-297

249

Table 12 Variation in product range and some leading properties of SMDS roduct

products (Sic et al., 1991)

Gas oil mode

Kerosene mode

Tops/naphtha

wt%

I5

25

Kerosene

wt%

25

50

Gas oil

wt%

60

Property

25

Gas oil (diesel)

Kerosene 150-250

Boiling range

“C

250-360

Density

kg/m3

780

Pour point

“C

-10

Cetane number

750

75

Smoke point

mm

Freezing point

“C

> 50 -47

meet all relevant specifications, as shown in Table 12 for typical product data (Sie et al., 1991). The smoke point of the kerosene and the cetane number of the gas oil are excellent; they can be used to upgrade low-quality liquids derived from thermal or catalytic cracking reactions. The theoretical maximum thermal efficiency for methane conversion to linear hydrocarbons (based on lower heating values) is 78%. The SMDS plant is heat-balanced so that no extra natural gas is used for utilities. The SMDS plant, if all goes well, will operate at 65% thermal efficiency, which is 80% of the theoretical maximum. An outline of the SMDS process configuration is shown in Fig. 26. Operations began in late 1993.

HUU

-1

c4

HZ @Naphlha

Natural gas d

SGP

synlm 9=

HPC + Distillation facifilies

-------Kerosene

FGasoil

SGP:

Shell Gsslllcallon Process

HPS:

Heavy Pmfffnprooess

HMU: Hydrogen Manvfacfullng Unit HPC:

Heavy ParefftnconversIon

HUG:

Hydrogenatlcm U~II

b%‘Pu: Wax ProductIon Uni1 Fig. 26. Simplified process flow scheme for SMDS.

250

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Fuel Processing Techmlo~y 48 (1996) 189-297

The first commercial SMDS plant, partly aided by its desirable location in Bintulu, Malaysia, is expected to cost about $660 million for a 12000 bpd output. It is expected that future plants will be up to 50000 bpd, with advantages of scale-up. As is the usual case, syngas manufacture constitutes over 50% of the total SMDS capital process cost. Future work will be directed to lowering this cost, improving the catalyst and the design of the synthesis reactors and general process integration. With further development leading to larger plants, it is expected that specific capital costs for remote gas areas could be in the range $25000-$30000 per daily barrel (Tijm et al., 1993). 4.6.6. Exxon advanced gas conversion technology Fully aware of the advantages of natural gas and of the promise of the use of available and remote natural gas, the Exxon Research and Engineering Company started work in 1980 on development of an advanced technology for converting the gas to high-quality refinery feedstock. They have developed the Exxon AGC-21 Advanced Gas Conversion Technology, consisting of an integrated three-step sequence of fluid bed syngas generation, a newly developed slurry-phase hydrocarbon synthesis and fixed bed upgrading of products (Eisenberg et al., 1993, Eisenberg et al., 1994). The syngas generation step incorporates a novel fluid bed reactor in which partial oxidation and steam reforming take place simultaneously in one reactor containing fluidized catalyst particles (FBSG). IT hydrocarbon synthesis (HPS) uses a proprietary high performance catalyst system (cobalt plus small amounts of a noble metal) combined with an advanced multiphase slurry reactor. The system presents significant challenges which have been met: improved heat removal, scale-up of the multiphase fluid reactor dynamics, maintenance of catalyst performance and separation of the liquid wax from the catalyst slurry. The final step involves mild isomerization and upgrading to water-white refinery feedstock, readily transported by pipeline or tanker. As in usual FT syntheses, the AGC-21 process yields products free of sulfur, nitrogen, vanadium, nickel, asphaltenes and polycyclic aromatics. Product options include premium-quality diesel and jet fuel, lube oils, waxes, and chemicals. The process provides flexibility in varying the nature of the FI product (Table 13). A simplified flow diagram of the Exxon AGC-21 process is shown in Fig. 27 (Eisenberg et al., 1993). synthesis 4.6.7. The Koelbel-Englehardt The Koelbel-Englehardt (K-E) synthesis is an offshoot or variant of the IT synthesis (Koelbel and Ralek, 1980). It is mentioned here because of its contribution to the enhancement of a growing interest in slurry-phase FT reactors. The process is suited

Table 13 Exxon process produces Product Naphtha Diesel/jet Cat feed Total

a flexible petroleum

product slate (Eisenberg

Max. cat feed 15 50 35 100

et al., 1994) Max. diesel/jet 30 70 0 100

1. Wet&r/Fuel

Fig. 27. Exxon AGC-21 1993).

gas conversion

Processing

process.

Technology

48 t/996)

Process development

189-297

units: flow diagram

251

(Eisenberg

et al.,

for operation with low hydrogen, CO-rich gas, as produced from industrial gases such as producer gas, blast furnace gas and gas from second generation coal gasifiers. The K-E process operated in 19X- 1953 at Meerbeck, Germany in the liquid three-phase column reactors found most suitable for this synthesis. The pilot unit had a production of 11.5 tonne per day of synthesis products which were essentially the same as those obtained in FT reactions. The reaction enthalpies of the ET and K-E reactions differ only by single or multiple values of the WGS reaction. It is worthwhile pointing out that Fischer and his co-workers were the first to use slurry reactors in the FT synthesis (Fischer et al., 1932). In general, this type of reactor uses finely divided catalysts suspended in a high-boiling oil (which may be a high-boiling Ff fraction). Hall et al. (1952) in England, Schlesinger et al. (1951) at the US Bureau of Mines, Sakai and Kunugi (1974) in Japan and Kuo (1983) at Mobil have all carried out work on slurry FT reactions. As mentioned earlier, Sasol is now operating a slurry bed FT reactor. On February 23, 1996, Sasol Limited and Haldor Tops&z of Denmark announced a technology cooperation agreement between the two companies. The agreement calls for combined promotion of Sasol’s slurry-phase distillate technology and Haldor Topsiie’s natural gas conversion technology (Oil and Gas Journal, 1996). Using this slurry-phase process, it was estimated by Sasol that a plant to produce a million barrels per day of diesel fuel would need about 85 billion cubic meters per year of natural gas feedstock; it could be built for $300-400 million. Combining Topsoe’s experience in natural gas conversion to syngas with Sasol’s slurry technology would allow economic conversion of any viable

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gas source worldwide to high-quality diesel fuel and potentially to petrochemicals. Plants would be built close to the gas field and converted straight from gas to middle distillates. Sasol envisaged that natural gas/diesel field plants could be build in modular fashion; additional units would be added as demand increased. Exxon has built and operated a large demonstration slurry unit in Baton Rouge, Louisiana. A number of firms, including Exxon and Shell, have participated with the DOE in the development of a slurry-phase IT reactor using precipitated iron catalysts at LaPorte, Texas (Rae and Gormley, 1990). 4.6.8. Rentech, Inc. Rentech, Inc. is a publicly traded Colorado corporation, which was operating a small, = 300-500 bpd, Fischer-Tropsch unit in Pueblo, Colorado. The plant used an iron based catalyst in a slurry reactor to produce naphtha, diesel fuel and wax. In March 1996, Rentech sold this plant to its Indian licensee, Donyi Polo Petrochemicals, Limited, Bombay, India. The unit was dismantled and shipped to Calcutta, India. It will be installed in Arunachal Pradesh, which is located in northeastern India. It will use as the feedstock approximately 100000 cubic meters per day of presently flared gas. It should be in operation using Rentech technology in early 1998 (Yakobson, 1996). 4.7. The Isosynthesis The Isosynthesis is part of the more generalized reaction systems associated with the FI process. This syngas reaction was developed during World War II by Pichler and Ziesecke (1949). Details of the project, actually started in 1941, were kept secret because its primary goal was the catalytic production of isobutane and isobutene, raw materials for high octane gasoline syntheses (Pichler, 1952). The Isosynthesis has been reviewed by Cohn (1956) and by Shah and Perrotta (1976). Syngas is catalytically converted predominantly to branched hydrocarbons using certain difficultly reducible oxides as catalysts. Development of the process was rapid but its commercial use was cut off by the successful development of new catalysts for the production of high octane gasoline from readily available petroleum. Although both use syngas as the feed, the Isosynthesis differs from the FI’S in several ways * The Isosynthesis gives high yields of isoparaffins rather than normal paraffins. * The catalysts are difficultly reducible oxides such as ThO, or ZIG, rather than reduced transition metals. * Isosynthesis temperatures and pressures are considerably higher than those used in the FIX. * Isosynthesis catalysts are not poisoned by sulfur to any great extent. * The distribution of products in the Isosynthesis differs considerably from those predicted by the ASF distribution (Table 14). - Optimum pressures for the production of liquid hydrocarbons in the Isosynthesis are between 30 and 60 MPa. At higher pressures, methane and dimethyl ether predominate. Thorium oxide (ThO,) is a good catalyst for the Isosynthesis reaction; the catalyst could be regenerated with air after several weeks so that it had considerable activity over

1. Wender/

Table 14 Comparison

of product distribution

Fuel Processing

for isosynthesis

Technology

48 (1996)

with that predicted

189-297

253

for ASF kinetics

Product selectivity/W% Isosynthesis

c, +cz C, +C, C 5+ CH,OCH, ‘I Conditions:

”

Predicted by ASF

17 43 37 3

17 22 61

45O”C, 600 atm, ThOz catalyst.

long times. Syngas is consumed in a 1.2CO/lH, ratio in a single pass. Although not affected by sulfur compounds, the synthesis should be carried out in chrome alloy steel or copper-lined vessels which exclude FTS catalysts (Shah and Perrotta, 1976). In addition to ThO,, both ZrO, and, to a lesser extent, CeO, are active Isosynthesis catalysts at 450°C and 150 atm. At temperatures below 375°C alcohols and dimethyl ether become the main products; above 5OO”C, low molecular weight hydrocarbons predominate. At 450°C 25% iso-C, compounds, chiefly isobutane, are formed, with about 16% of methane and 46% of liquid products containing mostly branched aliphatics plus some aromatics and naphthenes. The activity of ThO, is greatly increased by the addition of Al,O,; the best results were obtained with 20% of Al,O, added to the ThO, catalyst. Whereas Al,O, addition increased the formation of isobutane, addition of ZrO, to ThO, increased the formation of liquid products. Very little has been reported about the Isosynthesis process since the early work by Pichler and Ziesecke (19491, but interest in this reaction has been revived, chiefly because of the growing demand for isobutene and other branched hydrocarbons. Sofianos (1992) has reviewed the synthesis. The existing literature seems to reveal that the main products of the Isosynthesis reaction, namely isobutene and isobutane, can be obtained in sufficiently high yields only at high temperatures and pressures. The space time yields for the Isosynthesis are less than in the case of the synthesis of higher alcohols. The Isosynthesis reaction is not possible at low pressures as the formation of DME, lower alcohols and isobutanol predominates under these conditions. Isobutanol was one of the main products of Pichler and Ziesecke’s Isosynthesis reaction; this indicates a relationship between the Isosynthesis and the higher alcohol synthesis. Isobutanol plus other higher alcohols can be produced using a number of catalyst systems under milder conditions with greater yields than iso-C, compounds from the Isosynthesis. Large amounts of methanol are always present and there has come into play a driving force to realize a direct reaction between it and isobutanol or isobutene derived from isobutanol, to form MTBE. 4.8. Direct combustion for generation

of electrici

A technology that shows great oxygen-blown entrained gasification

promise for development involves pressurized of coal to produce syngas and its use in the

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Integrated Gasification Combined Cycle (IGCC) for the generation of electricity (Alpert, 1991). The US Department of Energy Clean Coal Technology program has been carrying out a program in this area using federal funds plus added private sector funds. A 100 MWe Cool Water Coal Gasification Combined Cycle Power test program, based on Texaco Coal Gasification, was completed in 1989. A team that included the Southern California Edison Company, Texaco Inc., The General Electric Company, Bechtel, a Japanese consortium led by Tokyo Electric Power Company, and EPRI as principals (ESEERCO and SOHIO Alternate Energy also were contributors) funded, built and operated the first integrated, commercial scale, coal gasification combined cycle power plant in the world (Spencer et al., 1982). Cool Water’s IGCC unit achieved net system heat rates of 9000 BTU per kWh with very low SO,r emissions of about 120 ppm and NO., emissions of about 25 ppm, with no particulate emissions. The Dow Chemical Company has been operating a second IGCC project at Plaquemine, Louisiana with a single gasifier producing fuel sufficient for 160 MW. The plant was originally fueled by natural gas; gasification of subbituminous coal to syngas was started in 1987. The Dow technology exceeded its rated design capacity and acceptable capacity by 1990. A commercial plant using this technology is planned by the Public Service Indiana Company. Several IGCC-type units are being operated or planned by US utilities and IGCC is being actively pursued or seriously considered in many countries. There is a great continuing interest in the commercialization of IGCC plants around the world and the reasons are not hard to find. Allowable emissions from electricity generation plants are becoming increasingly stringent and IGCC plants have superior environmental pet-formance. In IGCC operations, well over 99% of the sulfur is removed and NO, emissions are well below existing new source performance standard (NSPS) emission standards for new plants. The H,S and NH, in the gasification products are removed to a greater extent than is practical from combustion gas from conventional power plants. In IGCC, sulfur and nitrogen, and also particulates, are removed before combustion, rather than after combustion as in conventional pulverized coal combustion plants. The efficiency of IGCC plants is greater than that achieved in pulverized coal-fired plants. With further improvements in gas turbines, generation efficiencies using clean syngas could approach or exceed 45-55%. This is thus a conservation project, as more electricity is produced per ton of coal so that CO, emissions are reduced. It may be possible to capture and sequester the CO, produced. IGCC plants can be used in small or in large utilities, as they can be built in standardized modules. A flow diagram of an IGCC unit is shown in Fig. 28. There is a continuing development of improved gasifiers for IGCC. The principal gasifiers now in use are the Texaco, Dow (DESTEC), Royal Dutch/Shell and British Gas/Lurgi gasifiers, and coal gasification efficiency will be improved with ongoing developmental efforts. Another important proposal that expands the IGCC outlook is the possibility of co-producing electricity and chemicals derived from syngas. The gasifier can operate continuously and, if the requirement for generation of electricity falls, the syngas can be used, for example, for the synthesis of methanol, dimethyl ether, methyl formate or

1. Wemlrr/

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Technology 48 11996) 189-297

Fig. 28. Block flow diagram of Cool Water coal gasification feedwater heat exchange is not shown.

program

(Spencer

255

et al., 1986). The boiler

some other stable organic fluid which could be stored and used for peak shaving electricity generation. It may be feasible to synthesize methanol in a once-through operation (avoiding the cost of recycling), remove the methanol and use the unreacted syngas for electricity generation. The co-production of methanol and DME has been discussed earlier in this report (Hansen and Joensen, 1991; Brown et al., 1991). As mentioned earlier, Air Products and Chemicals, Inc. and Eastman Chemical are constructing an advanced coal-to-methanol plant to be located at the Eastman Chemical Company facility in Kingsport, TN. It is planned to demonstrate on a commercial scale the production of methanol from coal-derived syngas using the Liquid Phase Methanol, LPMEOH’” technology. Production of dimethyl ether (DME) as a mixed co-product with methanol for use as a storable fuel is also being considered. DME has a number of commercial uses. In a storable blend with methanol, the mixture can be used as a peaking fuel in IGCC electric power generating facilities; there are indications that the cost of energy so stored can be cheaper than using methanol alone. Also, in a storable blend with methanol, DME can be used to increase the vapor pressure of the mixture. The resulting higher volatility is expected to provide beneficial “cold start” properties to methanol being used as a diesel engine fuel. As indicated earlier, blends of DME and methanol can be used as a chemical feedstock or for the synthesis of newer oxygenated fuel additives. DME is also being viewed as an environmentally benign aerosol to replace Freon. Utilities are in the business of generating electricity. It is conceivable that, perhaps in cooperation with other manufacturers, utilities may eventually be able to broaden their horizons by entering or cooperating with the fuel and chemical markets via syngas. Routes from syngas to hydrogen, methanol, DME, ammonia, gasoline, diesel, jet fuels

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and other markets are well-established. IGCC electricity-generating plants may help satisfy the growing need for hydrogen gas, as aromatics in gasoline are restricted by the CAAA of 1990 and its amendments. Over five million tons of H, were produced in the United States in 1989 (McLamon and Cairns, 1989) and the demand for H, is growing rapidly. There is little doubt that the future of IGCC is bright throughout the world. The price of natural gas, which can be used in IGCC generation of electricity, is projected to rise. Coal is found throughout the world and plans for IGCC operation continue to be made. A group of Dutch utilities (SEP) is building a 250 MW IGCC unit in the Netherlands, scheduled to begin operation in 1993; it will use Shell gasification technology. The process is closely integrated and promises to have a high efficiency. Coal gasifiers generally yield syngas with high CO/H, ratios; interestingly, CO has a usable heating value about 18% higher than that of H, (Camell, 1977). Additional IGCC projects are planned to be built outside of the United States. Elcogas, a newly formed company headquartered in Spain, was created in 1992 with backing from the Commission of the European Communities THERMIE Programme. Major European electricity suppliers from Spain, France, Portugal, Italy, Britain and Germany will be involved in the construction and operation of this industrial plant. It will be situated at Puertollano, Spain and will be the largest IGCC plant in the world. Coal and petroleum coke will be gasified in this plant. Net performance is projected to be 45% and 335 MW of electrical power will be generated. The plant is scheduled to begin operation early in 1997. In the US, a 260 MWe IGCC plant is planned by the Tampa Electric Company. It will operate a Texaco pressurized, oxygen-blown, entrained-flow coal gasifier using 2000 tons of coal per day. The management, disposal and use of by-products from IGCC power generation have been discussed by Clarke (1991). A decade ago, when the Cool Water project was just getting started, IGCC was envisioned primarily as a coal-fired technology. Surprisingly enough, oil-fired IGCC seems perhaps closer to commercialization than coal-fired IGCC. Petroleum coke, by itself or mixed with coal, is also coming along for use in IGCC units. Oil refiners are finding that investment in IGCC technology could be a good capital investment because it uses low-value, high-sulfur fuels, such as bottom-of-the-barrel heavy oils, and petroleum coke as feedstocks while generating the hydrogen (from syngas) needed for desulfurization. IGCC also generates more than enough electricity for the production of steam for use in distillation (Lamarre, 1994). Italy is Europe’s largest consumer of oil for the generation of power. Most of the oil is high in sulfur (about 3%), for world petroleum reserves are becoming heavier and therefore higher in sulfur. A new European Community directive requires that the sulfur content of fuel oil be limited to 1% by 1998 and 0.25% by 2003. Currently, four Italian refineries plan to build IGCC units to begin operation between 1997 and 1999. This is a classic case of market forces (oil-based technology) overtaking the original plans based on coal based systems. The largest hurdle for IGCC technology use in US refineries is competition from natural gas, which is clean burning and could just as easily supply the electricity and

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Furl Prowssing

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48 (1996)

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257

hydrogen that coal, oil or coke gasification could supply. But IGCC remains an attractive option because it uses the growing supply of petroleum resid and petroleum coke for productive use. It is difficult to sign a long-term contract at a set price for natural gas; there is little doubt that its price will rise in the coming years. Mexico is also looking to the use of petroleum coke in IGCC units. China is interested in IGCC for the production of power from coal; concern about air quality is a driving force (Lamarre, 1994). Some IGCC advantages are listed below. - A clean environment. - Fuel flexibility. - High efficiency. + Low capital costs. - Modularity. * Phased construction. - Marketable by-products. * Co-products. - Ability to use “dirty” fuels. - Public acceptability. 4.9. Methanation,

syngas to substitute

natural gas (SNG)

The formation of methane from syngas was first reported by Sabatier and Senderens in 1902. A nickel catalyst is used almost exclusively to convert syngas to methane, although other transition metals may be used. The usual temperature range is 7001000°C. The dominant mechanism is dissociation of CO followed by rehydrogenation of the surface carbon atoms to methane. Adsorbed oxygen is removed from the catalyst surface as CO, by reaction with another CO molecule (Somorjai, 1981). Natural gas is a clean fuel in growing use throughout the world. Immense amounts of natural gas exist in remote areas of the world and new sources of methane are being found in all parts of the world. Lay (1993) estimated the remaining recoverable natural gas resource base of the lower 48 United States to be about 1.3 trillion (1.3 X IO’*) cubic feet at the start of 1993. Natural gas consumption is forecast to increase steadily at least through 2010 and gas supplies could be available at lower prices than in previous years. Increased demand comes from increasing use in electricity generation and also in industrial markets. Economics and environmental concerns are the main driving forces for increased use of methane; technological progress in locating and bringing natural gas to market are fueling these increased uses. There is only one plant in the United States that now converts syngas to substitute natural gas (SNG): the Great Plains Gasification facility in North Dakota. This plant was built with a loan guarantee of about $1.8 billion from the US Department of Energy. A consortium of four companies had contracted to purchase and distribute the SNG to their customers. The project has been a success from a technical point of view but oil and gas prices have not proved high enough to make the project economic. The SNG facility was turned over to the Basin Electric Cooperative to avoid US government losses. The facility is operated at a profit at this time; the consortium of gas

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I. Wender / Fuel Processing Technology 48 (I 996) 189-297

companies pays about $2.75-3.00 per million BTU of SNG, and the market price is approaching $2.00 for gas. The Dakota Gasification Company sponsored a working paper by MITRE to perform a techno-economic analysis of possible modifications of the Great Plains Gasification facility (Gray et al., 1991). The purpose of this study was to find ways of using all or part of the syngas available at the plant to produce liquid petroleum substitutes having a market value greater than the SNG currently produced, if current marketing agreements for SNG expire or otherwise become ineffective. Seven possible modifications were considered but there is no interest in implementation of any of them. The Great Plains Gasification facility is located at Beulah, North Dakota and produces SNG from North Dakota lignite using Lurgi dry-ash gasification technology (Sasol gave valuable assistance). The plant is a technical success and currently produces SNG at some 120% of design capacity. About 165 million standard cubic feet (SCF) of pipeline-quality gas are produced from = 17 000 tons of coal (lignite) per day. Current market prices of natural gas are lower than SNG production costs but the gas purchase contracts in effect enable the Dakota Gasification Company, owner of the facility, to operate the plant economically at the time of writing. 4.10. Fuel cells Fuels cells are included in this report because the primary reactants are hydrogen, syngas and/or carbon monoxide. A review of fuel cell technology has been given by Kinoshita and Cairns (1984). A thorough overview of fuel cell technology is available in the Proceedings of the International Conference on Fuel Cells (Wedaa, 1994). Although usage of these syngas components is small at present, fuel cell development is in a growth stage with a promising future. Fuels cells have the advantage of converting energy from a chemical reaction directly to electricity. Although fuel cells have been investigated for well over a century, NASA, in the 1960s furnished a new emphasis for their development for space applications. They worked well in space but were too costly for general use on earth. Fuel cells generate electricity without combustion. They differ from batteries, which are energy storage devices; fuel cells are energy conversion devices. Reactants are supplied continuously in a fuel cell. The great advantage of the fuel cell is that it is free of ¬-type thermodynamic limitations. Some advantages of the production of electricity from fuel cells are: no moving parts, no frictional wear, virtual absence of noise, minimal environmental pollutants, efficiency independent of size, and modularity that allows installation in almost any area so that power line investment and transmission losses are decreased. But there are disadvantages, including reliability, overall high costs, actual efficiency, sluggish reaction kinetics and ohmic losses (Prentice, 1984). Recent research has focussed on overcoming the two primary impediments to the widespread use of fuel cells: high initial cost and short operational lifetimes. The variety of fuel cells may be classified by the type of electrolyte used. In order of increasing temperature, they are: the polymer electrolyte fuel cell (PEFC) at 80°C the alkaline fuel cell (AFC) at 80-9O”C, the phosphoric acid fuel cell (PAFC) at 260°C the

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259

Y

GENERATION WITHOUT COMBUSTION Fuel cells are like batteries that run on fuel. Oxidation of hydrogen and carbon monoxide at the anode releases electrons to an external circuit, which conducts them to the cathode, where they combine with oxygen atoms in a reduction reaction. The negatively charged oxygen ions thus produced are transported through an electrolyte to the anode. Solid-oxide fuel cells have a ceramic electrolyte and operate at temperatures high enough for internal reforming of methane fuel into the hydrogen and carbon monoxide needed for the oxidation reaction.

Fig. 29. Generation

of electricity

i’l

Fuel (CH,

Jntema1 refomling 3H1+ CO

)

CathCdC

without combustion:

operation

of a fuel cell (Douglas,

1994).

molten carbonate fuel cell (MCFC) at 650°C and the solid oxide fuel cell (SOFC) at 1000°C. All are of interest, with the PAFC, MCFC and SOFC gaining most attention. SOFC units do not have the nuisance of using liquid electrolytes and are of comparatively simple design, involving inexpensive ceramic or metallic components. They have the potential of long service lives: nearly 10 years, or almost twice that projected for cells with liquid electrolytes. EPRI, Westinghouse and other companies are actively supporting research on SOFCs with the aim of achieving commercial viability for dispersed electrical power generation by the year 2000 (Douglas, 1994). The mode of generation of a fuel cell is shown in Fig. 29. Electrons are released to an external circuit by oxidation of H, and CO at the anode. They are conducted via an external circuit to the cathode where they combine with oxygen atoms in a reduction reaction. The negatively charged oxygen ions are transported through an electrolyte to the anode. SOFCs with a ceramic electrolyte operate at high temperatures (1000°C) so that internal reforming of methane to H, and CO for the oxidation reaction can take place (Douglas, 1994). Challenges facing the commercialization of fuel cells have been outlined by delegates from 23 countries who met recently at the Commonwealth Institute, London (Fourth Grove Fuel Cell Symposium, 1995).

5. Category 3. The hydroformylation

(0x0) reaction

5.1. Introduction The hydroformylation (0x0) reaction consists of the reaction of an olefin with syngas to form aldehydes or alcohols with one more carbon that the starting olefin. The reaction

I. Weruler / Fuel Processing Technology 48 (1996) 189-297

260

was discovered by Roelen (1938), Roelen (1943) while he was studying the effect of added olefins on the Fischer-Tropsch reaction using a heterogeneous cobalt catalyst. The hydroformylation reaction was first conducted in Germany as if it were a heterogeneously catalyzed reaction, but it was found necessary to continuously add cobalt salts to the reactor. Roelen recognized that the reaction was catalyzed by soluble cobalt carbonyls and that the hydroformylation is a homogeneously catalyzed reaction. It was the first industrially important reaction catalyzed by soluble complexes; this oldest homogeneously catalyzed reaction is still the most important of its kind. Hydroformylation of olefins forms the basis of a worldwide industry which accounts for over 7 million tons of products per year (Papp and Baerns, 1991). There are a number of reviews of the hydroformylation reaction (Pino et al., 1977a; Pruett, 1981; Keim, 1983; Henrici-Olive and Olive, 1984; Whyman, 1985; Parshall and Ittel, 1992). The chief uses of syngas are based on the use of traditional heterogeneous catalysts: the manufacture of hydrogen gas, the synthesis of methanol and the FT reaction. Hydroformylation is the fourth largest use of syngas, utilizing over 200 billion SCF per year. The simplest hydroformylation reaction takes place with ethylene Co*(CO), CH,CH,CHO +

CH,=CH,+CO+H2 or

(5.1)

Rh complexes

Hydroformylation and many other reactions catalyzed by soluble transition metal complexes are known; they yield high-value chemicals. In these syntheses, the reactants, catalysts and products are present in the same phase, usually the liquid phase. Heterogeneously catalyzed reactions take place on metals usually supported on various oxides; homogeneous transition metal catalysts involve soluble complexes at low temperatures (< 250°C) with high selectivity. In contrast to the complex reactions that occur on surfaces in heterogeneously catalyzed reactions, the mechanisms of homogeneously catalyzed reactions are reasonably well understood. Such catalysts have significant industrial applications and will find increased use in the future. The hydroformylation of olefins with syngas is a rapid reaction catalyzed chiefly by soluble complexes of cobalt or rhodium. The reactivity of various transition metal catalysts, compared with cobalt, follows the general pattern Rh 103-

104

> Co > Ru, Ir > Mn > Fe 1

10-Z

1O-4

10-h

Although the hydrogenation of olefins and of carbon monoxide to methane and water is thermodynamically more favored than hydroformylation, these reactions are suppressed or diminished owing to the high selectivity of the 0x0 catalytic system. The hydroformylation of propylene to form C, aldehydes and alcohols is the basis for an important industrial application of the hydroformylation reaction. CH,CH = CH, + CO + H, “zRhCH,CH,CH,CHO

+ CH,CH(CH,)CHO

The aldehydes can be converted to alcohols during the reaction or separately over a hydrogenation catalyst. The synthesis of n-butyraldehyde from propylene accounts for over 4 million tons per year of 0x0 products.

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The most important application of the 0x0 reaction hexanol, which is widely used as a plasticizer alcohol

is in the synthesis

base

2CH,CH,CH,CHO

CH,CH,CH,CH

261

+ CH,CH,CH,CH(OH)CH(CH,CH,)CH)CHO

of Z-ethyl-

-H,O

-+

(5.3)

= C(CH,CH,)CHO?

CH,CH,CH,CH,CH(CH,CHs)CH,OH

(5.4)

The hydroformylation reaction is applied to a large number of olefins, including unsaturated oils, fats, polymers, terpenes, carbohydrates and even to steroids. In addition to plasticizers, products include solvents, synthetic detergents, flavorings, perfumes, and products for the fine chemical industry in general (Pino et al., 1977b). 5.2. Catalysis by cobalt Cobalt metal and most cobalt salts form cobalt carbonyl, Co,(CO),, under hydroformylation conditions (200-300 atm, 1: 1 HJCO, and 1 IO-200°C). Catalyst concentrations are about 0.1 - 1% of metal to olefin. Ratios of n-butyraldehyde to isobutyraldelyde (n/iso ratios) of about 4: 1 or higher can be achieved. The straight-chain compound is the most desirable product. About 2% of olefin hydrogenation occurs and some 5% of high-boiling products are generally found. The cobalt catalyst functions as both a double-bond isomerization and an 0x0 catalyst. Although internal olefins are thermodynamically favored over terminal olefins, good yields of straight-chain aldehydes and alcohols can be obtained from internal olefins with cobalt catalysts; the double bond migrates to the terminal position during the reaction. The normal isomer is favored at lower temperatures and higher pressures of carbon monoxide. The cobalt-catalyzed hydrofotmylation reaction was operated successfully for over 20 years with little technical change, but the need to improve the cobalt-catalyzed process finally became obvious. There were apparent demands to conduct the reaction at lower pressures (the high pressure is necessary to stabilize the cobalt carbonyl intermediates), to obtain higher n/iso ratios. to improve catalyst stability and recovery and to cut down on side reactions. In the late 1950s it was found that replacement of the carbonyl groups in Co,(CO), by organophosphines and similar ligands formed complexes of higher thermal stability. This led to two new hydrofotmylation processes, the Shell phosphine-modified cobalt process (Slaugh and Mullineaux, 1968; Tucci, 1970) and the Union Carbide phosphinemodified rhodium process (Fowler et al., 1976). In the Shell process, modified cobalt carbonyls such as HCo(CO),(PBu,) are used. This and similar cobalt carbonyl complexes, although less active than Co,(CO), even at 180°C are much more active for hydrogenation to yield n/iso alcohol ratios of = 7: 1 (Falbe, 1970). The tributylphosphine-cobalt catalyst is more stable and can be used at about 100 atm rather than 200-300 atm with the cobalt carbonyl system (Parshall and Ittel, 1992). The alcohols can be obtained by distillation and the catalyst recovered.

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5.3. Catalysis by rhodium

Although rhodium carbonyls are 103-lo4 times more active than cobalt carbonyls, they give mostly branched aldehydes, i.e., the n/iso ratios are very low. But active rhodium catalysts with excellent selectively to linear aldehydes (high n/iso ratios) were obtained with PBu, or P(C,H,), ligands attached to rhodium (Pruett and Smith, 1969, 1970). Union Carbide started commercial production of these modified rhodium catalysts in 1976 (Fowler et al., 1976, Brewester, 1976); n-butanol yields of about 85% can be obtained. A comparison of the phosphine-modified cobalt process and the phosphine-modified rhodium process is given in Table 15. Hydroformylation is a versatile reaction in which a -CHO or -CH,OH group is added to the carbon chain of an olefin. It has a myriad of applications, a few of which are given below. With large amounts of industrial olefins or if a particular chemical is desired, there is the choice of three complementary hydroformylation routes which compete for use. The choice, use of straight cobalt or phosphine-modified cobalt or rhodium, depends on the particular olefin and the desired product. Phosphine-modified rhodium is the clear choice to produce 2-ethylhexanol from propylene plus syngas. For some applications, the latter choice is not versatile enough and the non-liganded cobalt would be used for hydroformylation of mixed olefins of higher carbon number. Phosphine-modified cobalt is best for the production of higher alcohols rather than aldehydes. The choice is not always clear-cut. It is worth noting that many hydroformylation plants worldwide still use the original cobalt catalyst with no added ligands, but the use of rhodium complexes as hydroformylation catalysts is growing rapidly. Because this technology is so versatile and has such different applications, a few more examples of specific commercial hydroformylation uses will be given. Vitamin A is produced on a relatively large scale and competition is keen in this area. Both Hoffmann-La Roche and BASF employ hydroformylation as a key step in the commercial production of vitamin A (Parshall and Ittel, 1992). A commercial route to butanediol based upon the hydroformylation of ally1 alcohol has been developed. Shell manufactures a number of unsaturated aldehydes based on the selective hydrofotmylation of unsaturated intermediates derived from their metathesis plant (Parshall and Ittel, 1992). The products are used in perfumes and in other commercial products. Ajinomoto manufactured monosodium glutamate by the hydroformylation of acrylonitrile for a

Table 15 Comparison of phosphine-modified cobalt and rhodium processes Catalyst precursor Phosphine/metal ratio PIEssure Temperature Catalyst concentration (% metal/olefm) n/is0 ratio Olefin hydrogenation High boiling products

Co,(CO), + n-Bu,P 2:1 50-100 abn 160-200°c 0.6 7:1 10% 1%

HRh(COXPPh 3)3 + PPh 3 50-loo:1 l-25 atm 60- 120°C 0.01-0.1 8:1-16:l 5-10% low

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decade ending in 1973 (Yoshida, 1978). Shell carried out the hydroformylation of various internal olefins to convert them to terminal alcohols There are numerous examples of the use of the hydroformylation reaction. There are continuing advances in hydrofotmylation catalysis and the area is still the subject of active investigation. Ruhrchemie/Rh&te-Poulenc has commercialized a propylene hydroformylation process using a trisulfonated triphenylphosphine rhodium ligand (Kuntz, 19871. This complex is soluble in water but insoluble in the organic phase, so that essentially no losses of rhodium occur to the organic phase. With water-soluble, ligand-modified rhodium catalysts, products remain in the organic phase and modified catalysts are immobilized in the water phase. The metal catalyst is prevented from leaching into the process products. The catalysts are separated by decantation from unreacted olefins and products, greatly simplifying the process as compared with other 0x0 process routes. This variant of the 0x0 process is in commercial use (Comils and Weibus, 1995) and has great potential for extension to other carbonylation reactions. It is necessary, of course, to recover the rhodium essentially completely in hydroformylation reactions. Even cobalt should be recovered. 5.4. Chiral catalysis An sp3 carbon atom with four different groups attached to it can exist in two enantiomeric forms. Each enantiomer is a mirror image of the other. Most biologically active compounds contain at least one optically active carbon; sugars and proteins in living organisms are found in only one of the possible optically active forms. The addition of chiral ligands to cobalt or rhodium catalysts in the hydroformylation reaction can yield products in which one enantiomer exists in excess (Consiglio et al., 1973; Watanabe et al., 1979; Botteghi et al., 1980). This is an exciting field of research that promises to grow in importance because of its environmental and conservation benefits. As examples, most pesticides are optically active but often only one of the emutiomers is biologically active. We can synthesize the active enantiomer and spray fields with half the amount of pesticide now in use. The inactive enantiomer at present has no function but to contaminate the area. The drug thalidomide, whose use was found to be catastrophic, exists as a racemate, an equal mixture of two enantiomers. Only one of the two enantiomers is toxic; the other appears to be useful in other safe treatments. These optically active products are usually made using selective homogeneous catalysts, although heterogeneous catalysts are under development. Chiral syntheses can be used to manufacture high-value pharmaceuticals, herbicides, insecticides, fungicides, etc. A well-known example, not made via the hydroformylation reaction, is the manufacture of L-dopa, which is used in the treatment of Parkinson’s disease, via a homogeneous catalytic asymmetric hydrogenation using a chiral rhodium complex (Knowles, 1983). 5.5. Mechanism

of the hydroformylation

reaction

Although there are differences in the mechanistic routes involved in the cobalt, cobalt-phosphine and rhodium-phosphine processes, the steps in the generally

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“accepted” reaction routes are similar. It is important that the first step involves the removal of CO from the added catalyst complex to give a coordinatively unsaturated species (16-electron species). The olefin then adds to the vacated orbital, followed by insertion of the coordinated olefin into the Co-H or Rh-H bond (Orchin and Rupilius, 1972). cc+ (CO), + H, + HCo(CO), + R’CH=CH,

ZHCo(CO)~

HCo(CO), + CO

+ HCo(CO),

R’CH? CH, e HCo(CO),

RICH+ CH, HCo(C0) 3

RCo(CO), +m T-- RCo(CO),

RCo(CO), P

RCOCo(CO),

RCOCo(CO),

wcoco(co), -

HZ RCOCo(CO),

+

I.--

RCHO + HCo(CO),

__ RCHO + Co,(CO), HCo(CO),

R = R’CH,CH, or R’CHCH,

The aldehydes can be reduced to alcohols in the same system or they may be hydrogenated to alcohols in a second step. Similar mechanistic sequences can be written for the rhodium-catalyzed hydroformylation of olefins.

6. Category 4. The Mobil methanol to gasoline (MTG) and related processes 6.1. Introduction: zeolites

Although invented by scientists at Mobil over 15 years ago, the Mobil MTG process is still considered a breakthrough in catalytic science (Meisel, 1981; Chang, 1983; Chen et al., 1989; Tabak and Yurchak, 1990). It has been hailed as the first new route for the production of gasoline in over 40 years. The process is based on the conversion of syngas to methanol or dimethyl ether (formed by dehydration of two molecules of methanol) over a synthetic pentasil zeolite family, ZSM-5, invented by Mobil in the 1960s (Argauer and Landolt, 1972). ZSM-5, the key element in the MTG process, has a unique pore structure and unusual catalytic properties. It has two sets of intersecting channels with openings of = 6 A. One set consists of elliptical lo-membered ring straight channels; the other is a set of tortuous sinusoidal channels (Fig. 30). The channel opening is just large enough to allow

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Fig. 30. Structure of ZSM-5.

molecules with 10 or fewer carbons to wend their way through the crystalline structure. If the molecule is small, it moves rapidly through the zeolite and experiences less cracking. Molecules containing more than 10 carbon atoms are unable to leave the catalyst pores and eventually reform until they are of the right size to emerge from the catalyst. A word about zeolites, which have become possibly the most useful catalysts in chemical processing with a rapidly growing potential for new applications (Chen et al., 1989, Haag et al., 1987). Zeolites are crystalline, three-dimensional aluminosilicates in which the building blocks are silica and alumina tetrahedra. The framework will have a net negative charge because every oxygen in the “infinite” lattice is shared by two tetrahedra, with silicon being tetravalent but aluminum only trivalent. The negative charge is balanced by exchangeable cations: a general structure may be written as (Chang, 1983)

M,>,,(AlO,1,WO, >y. x&O where n is the charge on the cation and x is the water of hydration. The SiOJAlO, ratio varies in different zeolites; it is high in ZSM-5. The mouth of a particular zeolite channel consists of a ring of a fixed number of tetrahedra, which determines the diameter of the mouth or the pore size. Over 30 natural zeolites are known and many more than that number have been synthesized in the laboratory. Although the central atoms of all natural zeolites are dominated by Si and Al, chemically related atoms such as B, P, Ge, Ga, etc. can be incorporated into zeolites. 6.2. The MTG reaction Molecular shape selectivity in these aluminosilicates was first reported by Weisz and Frillette (1960). Weisz et al. (1962) defined two types of shape selectivity: reactant selectivity, where certain molecules enter into the zeolite and others are excluded by virtue of their shape and size, and product selectivity, where some of the products formed within the pores are too bulky to diffuse out. These latter molecules are either

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converted to smaller molecules or eventually deactivate the catalyst by blocking the pores. Csicsery (1976) has reviewed the modes of action of shape selective catalysts. The Mobil (MTG) process may be represented by the following equations 2CH,OH e CH,OCH, + H,O

(6.1)

-H,O

CH ,OCH s +

C 2 - C s olefins

(6.2)

C, - C, olefins + paraffins, cycloparaffins, aromatics

(6.3)

Selectivity to gasoline-range hydrocarbons is more than 85% with virtually no compounds heavier than C,, being formed. Liquefied petroleum gas (LPG) makes up most of the remainder of the product. One hundred tons of methanol is converted to nearly 44 tons of hydrocarbons and 56 tons of water. The hydrocarbons produced contain 95% of the energy in the methanol feedstock; the exothermic heat of reaction contains the remaining 5% of energy. The first (and only) commercial MTG plant came on stream in 1985 in New Zealand, originally producing 14500 bpd of gasoline from natural gas. Approximate distributions of products obtained from Sasol’s fixed bed ARGE and fluid bed (Synthol) reactors are compared with those obtained from Mobil’s MTG process in Table 16 (Dry, 1981; Mills, 1977; Roper, 1983). A normalized distribution of aromatics obtained in the MTG process is, in wt%: benzene 4.1, toluene 25.6, ethylbenzene 1.9, o-xylene 9.0, m-xylene 22.8, p-xylene 10.8, trimethyl substituted benzenes 14.1, other aromatics 12.4 (Walker, 1985). Essentially three steps are involved in the MTG process: conversion of methanol to dimethyl ether (DME), initial formation of a carbon-carbon bond, and finally aromatization with hydrogen transfer. Crude methanol containing 17% water may be used in this process. Commercial development of the MTG process is carried out in a fixed bed reactor in New Zealand. Chang (1983) has reviewed the chemistry and operating characteristics of the process. In 1984, a 100 barrel per day demonstration plant was built in Wesseling,

Table 16 Product distributions

from Sasol (FT) and Mobil (MT@ processes

Temperature/K Pressure/atm Feed Product distribution Light gas C, -C, LPG c,-c, Gasoline C,-C,, c,3-Cl9

Heavy oil, C ,9+ Oxygenated compounds Aromatics, % of gasoline

ARGE (fixed bed) FT

Synthol (fluid bed) m

Mobil (fixed bed) MTG

490-520 26 1.7H,:lCO

633-685 22 3H,:lCO

63-685 14-24 CH,OH

11.0 11.0 25.4 14.0 37.0 2.3 0

20.1 23.0 39.0 5.0 6.0 7.0 5

1.3 17.8 80.9 0 0 0 38.6

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near Bonn, Germany (Tabak and Yurchak, 19901. The major concern in both types of reactor is in the control and dissipation of heat generated by the exothermic reactions involved. A process flow diagram of the New Zealand MTG plant is shown in Fig. 31 (Tabak and Yurchak, 1990). Methanol is first dehydrated to an equilibrium mixture of methanol/DME/water which releases about 20% of the overall heat of reaction. The mixture is then fed into reactors containing the ZSM-5 catalyst to convert both the methanol and the DME to gasoline and water. The raw gasoline is distilled to reduce the content of durene (1,2,4,5tetramethylbenzene) to a satisfactory level. The gasoline produced is of high quality with a RON (research octane number) of 92.0-92.5. The mechanism of the conversion of methanol and DME to gasoline is still controversial and is discussed elsewhere (Chang, 1988; Hutchings and Hunter, 1990). The 100 barrel per day second generation MTG fluid bed reactor built and operated in Germany gave higher yields of gasoline than the fixed bed MTG reactor. With improved energy utilization and advantages in yield and quality of the products, there is a cost advantage of at least 10% for the fluid bed semiworks (Mills, 1993). This successful demonstration of the MTG fluid bed operation is available for commercialization but, with the availability of cheap petroleum, there is no interest in further pursuit of the MTG process at this time. The MTG plant was the result of an innovative discovery and concept. However, the cost of producing high-quality gasoline from natural gas via syngas has proved too expensive for the New Zealand government, which originally planned to have an “indigenous” source of gasoline and to lower their balance of payments in this way. The MTG plant in New Zealand has been sold to a private company which sells both methanol and high octane gasoline, the relative amounts changing with market prices and demand. Compared with related processes, Sasol would seem to be at a disadvantage for several reasons: coal is the syngas source, a significant amount of methane produced has to be reconverted to syngas, and the IT product consists of a multiplicity of hydrocarbons and oxygenates. Sasol built new and improved reactors and developed sophisticated and complicated processes to separate the daunting mixture of products so as to market everything possible. In 1994-95, fuels accounted for 44% of Sasol’s production with chemicals accounting for another 31% (mining, fertilizers and explosives constitute

Pro&Xx to

Trertmonr

Fig. 3 I. Schematic

of New Zealand gas to gasoline complex.

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268

the remainder). Sale of FT waxes has turned out to be a profitable business. Sasol has not hesitated to acquire technology from abroad and competes on the world market. The MTG process in New Zealand produces only methanol and gasoline; the methanol is traded on the world market. 6.3. MTE, MT0 and MOGD reactions Mobil has also developed several processes based on methanol as a source of ethylene (MTE), of olefins (MTO) and of gasoline plus diesel (MOGD) (Tabak and Yurchak, 1990). These are promising catalytic routes to valuable products from methanol. By reducing the zeolite acidity and raising the operating temperature to > 5OO”C,Mobil succeeded in modifying the rate of olefin formation so that over 80% of the product consists of C,-C, olefins. This is known as the MT0 process. Oligomerization, disproportionation and aromatization of the C 2-C l0 olefins formed in the MT0 process are the basis for Mobil’s MOGD (olefin to gasoline and distillate) process (Garwood, 1983). Premium-quality distillate and low-pour-point stable jet fuels that meet all commercial and military specifications are produced. These processes have not come into commercial practice as yet but they have considerable promise for the future. 6.4. Topside integrated gasoline synthesis (TIGAS) The TIGAS process, developed by Haldor Topsae A/S, has the unique feature of integrating the methanol synthesis with Mobil’s MTG process into a single loop without isolation of methanol as an intermediate. The process combines steam reforming and autothermal reforming for the production of syngas (Topp-Jorgensen and Rostrup-Nielsen, 1986, Topp-Jorgensen, 1988). The MTG process takes place in three sequential steps: production of syngas, synthesis of methanol, and conversion of methanol to hydrocarbons over a ZSM-5 catalyst. However, these syntheses are preferably carried out at different pressures. Syngas is typically manufactured at 15-20 atm; the methanol loop normally operates at 50-100 atm and the MTG fixed bed process is carried out at 15-25 atm. The TIGAS process aims to modify operating catalysts and conditions so that the pressures level out. This process, as shown in Fig. 32, combines steam reforming with autothermal reforming using oxygen in a secondary reforming step for syngas production. This allows syngas production at a pressure level near that of the methanol synthesis.

Recycle Gas

7

Natural GassA Oxygenate Steam

*

sWG??s Production

&

Gasoline -

Synthesis

Synthesis

Fig. 32. Gasoline from natural gas with the TEAS

process (Topp-Jorgensen,

1988).

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TIGAS uses a multifunctional catalyst based on a combination of the synthesis of methanol, which is an equilibrium limited reaction, and the dehydration of methanol to DME, which is an equilibrium unlimited reaction (Lee et al., 1992). The methanol synthesis takes place over a copper-containing catalyst and the dehydration takes place over an a-A1203 catalyst. The use of a combined methanol/DME synthesis in the TIGAS process using a multifunctional catalyst system helps to allow a low operating pressure for the integrated process. A relatively large amount of hydrogen in the gasoline synthesis significantly reduces the olefin content of the gasoline product. A semi-industrial pilot plant using the TIGAS process was built and operated in Houston, to a production of Texas. The capacity of the plant was 400 Nm3 hh ‘, corresponding MTPD of gasoline. A process unit was operated for IO 000 hours during the l984- 1987 period. The main advantages of the TIGAS process are process flexibility and lower investment (Mills, 1993).

7. Category

5. Methanol

plus syngas or CO for synthesis

of chemicals

7. I. Introduction A classification of syngas chemistry was given earlier. Category 1 dealt with syngas as a source of hydrogen, which can be considered the largest use of syngas. Category 2 dealt with the direct conversion of syngas to fuels and chemicals; this encompassed the direct use of syngas for conversion to methanol, in the Fischer-Tropsch reaction, in the direct combustion of syngas for the generation of heat and electricity, in the synthesis of methane, and in the direct synthesis of higher alcohols. The fourth largest use of syngas after hydrogen, methanol, and FT products is in the hydroformylation (0x0) reaction, the conversion of olefins and syngas to aldehydes and alcohols containing one more -CHO or -CH,OH group than the starting olefins. This use of syngas is dealt with early because of its importance and because it does not fit well into other categories. Category 4 dealt with the conversion of methanol to fuels and chemicals, including the Mobil MTG, MTE, MTO, and MOGD reactions. It seems surprising that there is no industrially practiced direct conversion of syngas to a major chemical, hence there is no category designation for chemicals made directly from syngas. This absence of such a category is not due to a lack of efforts to synthesize major chemicals, such as ethylene glycol. acetic acid, acetic anhydride, acetaldehyde and ethanol, among others, directly from syngas. The theoretical reaction stoichiometry for the conversion of some of these chemicals and the Hz/CO ratio of the syngas required are listed in Table 17. 7.2. On the direct synthesis of chemicals from syngas It would seem highly desirable to synthesize cially ethylene glycol, acetic acid and acetic

chemicals directly from syngas, espeanhydride. The chemical industry is

270 Table 17 Theoretical 1983) a

I. Wender / Fuel Processing

reaction stoichiometry

for some important

Chemical

Reaction

stoichiometry

acetaldehyde acetic acid acetic anhydride ethylene glycol ethanol ethylene

3H, +2CO+CH,CHO+H,O 2H 2 + 2C0 --f CH ,COOH CH ,OAc + CO + Ac,O 3H, +2CO+ HOCH,CH,OH 4H, +2CO+CH,CH,OH+H,O 4H, +2CO+H,C=CH,

Technology 48 (1996) 189-297

chemicals

if made directly from syngas (Aquil6 et al.,

H, /CO ratio in syngas 1.5 I 0 0 2 2

AC is CH,CO. a Reprinted by permission from Hydrocarbon by Gulf Publishing Co., all rights reserved.

lb feed per lb product 1.4

1 I 1.5 1.4 2.3 Processin g, March 1983, p. 57, copyright

1983

increasingly being subjected to pressures to minimize or eliminate waste in the manufacture of products (Sheldon, 1992, Sheldon, 1993). Although efforts to synthesize several chemicals directly from syngas have not been successful as yet, their direct synthesis is an extremely worthwhile goal. It is not well perceived that the main source of waste (defined as everything but the desired product) in the synthesis of organic chemicals is inorganic salts usually formed in acid-base neutralizations, i.e. sodium chloride, sodium sulfate, ammonium sulfate and other salts. We shall see later that some direct carbonylation reactions, including the catalytic carbonylation of methanol to acetic acid, the carbonylation of propyne (methylacetylene) in the manufacture of methyl methacrylate and the synthesis of the drug ibuprofen via a catalytic carbonylation, are elegant examples of processes with high atom utilization (Sheldon, 1994). It is revealing to mention, as examples, efforts made by the Japanese to synthesize four of these chemicals directly from syngas; similar efforts have been made in laboratories throughout the world with the same goals. The Japanese C-l Chemistry National Project started in December 1980 and was officially concluded in July 1988 (Nakamura, 1990). Fourteen chemical companies, one private research institute and the National Chemical Laboratory were involved in the pursuit of the synthesis of four chemicals directly from syngas: acetic acid, ethylene glycol, ethanol and ethylene (Research Association for Cl Chemistry, 1989). After a thorough search for catalysts, bench-scale units were developed and pilot tests were designed but were subsequently cancelled. None of the direct syngas routes to these chemicals is now used industrially, although engineering data are available for the resumption of efforts toward commercialization. The Union Carbide Corporation patented a process for the direct production of C, oxygenated compounds from syngas using rhodium catalysts (Bhasin, 1975, Bhasin et al., 1978). However, these and other researchers in several laboratories found that acetic acid could not be produced selectively from syngas at the desired rate with rhodium catalysts or modified rhodium catalysts (Ichikawa, 1982; Watson and Somorjai, 1981; Orita et al., 1984; Arakawa, 1984; van der Lee et al., 1986). Other Group VIII catalysts

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such as cobalt, ruthenium or palladium gave poor conversions and selectivities for acetic acid. The direct synthesis of ethylene glycol from syngas is highly desirable, with seemingly complete conversion of feed to product. The Union Carbide Corporation pioneered work in this synthesis, discovering a homogeneous rhodium catalyst that seemed effective for this reaction (Dombeck, 1983). However, catalyst activity was low, the required syngas pressure was too high and selectivity was too low. Further work showed that bulky alkylphosphine-modified rhodium catalysts improved the results but conversions and selectivity were too low under the severe reaction conditions. Detailed engineering of a three ton per day pilot plant for the direct conversion of syngas to ethylene glycol was completed (Nakamura, 1990). But ethylene glycol is today not made directly from syngas, although potentially promisin g routes to the glycol involving the use of syngas are given in Fig. 33. Attempts to synthesize ethanol directly from syngas were also uneconomical but yielded data for future work (Pruett and Walker, 1974; Rathke and Feder, 1978; Williamson and Kobylinski, 1979; Kiennemann et al., 1987). The reaction of methanol with syngas to yield ethanol (the homologation of methanol) will be discussed later. 7.3. Scope of methanol-syngas

chemistry

Although the introduction of a one-step synthesis of chemicals from syngas is an attractive concept, it has been demonstrated that many industrial chemicals are produced

Ml3TllANOl.

Cfl3C0011

Methyl

formate

ffCOOCHs

Fig. 33. Methanol based route to chemicals

via syngas.

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more economically by multi-step (sequential) reactions; each step is carried out at conditions that are favorable thermodynamically and kinetically. But limitations of the selectivity or activity of direct routes to chemicals from syngas have been circumvented by the use of methanol as a feedstock. A look at the special nature of methanol is in order (Wender, 1984). It is made commercially and economically from syngas in over 99% selectivity. It is especially susceptible to nucleophilic attack in the presence of suitable acids because of the lack of steric hindrance on a one-carbon alcohol; methanol is over 100 times more reactive than ethanol in this type of reaction. The carbon-oxygen bond in methanol is the strongest C-O bond of any alcohol, with only ethanol having a comparable bond energy. Methanol and transition metal complexes react with HI or I, to form CH,I as a reactive intermediate in many syntheses. Methyl-metal bonds are readily formed with transition metal complexes. The CH, group readily undergoes a migratory insertion reaction to form acyl-metal (CH,CO-M) bonds. Methyl alcohol is a stronger acid (loses a proton more easily) than any other alcohol. The methoxide ion CH,O- is strong enough to react with the weakly basic CO to form methyl formate, which has a chemistry of its own. There are excellent reviews of the industrial synthesis of major chemicals from methanol (Parshall and Ittel, 1992; Agreda and Zoeller, 1993; Fahey, 1987; Keim, 1983; Papp and Baems, 1991; Herman, 1984). Some examples of products potentially synthesized from either methanol or formaldehyde are shown in Fig. 33. Formaldehyde was the main chemical derived from methanol, although the amount of methanol devoted to the synthesis of methyl t-butyl ether has surpassed formaldehyde use. The reactions of methanol may be divided into simple carbonylation (methanol plus CO>, reductive carbonylation (methanol plus syngas) or oxidative carbonylation (methanol plus CO and oxygen). Syntheses from formaldehyde are divided similarly: carbonylation (formaldehyde plus CO plus either H,O or ROH) or reductive carbonylation (formaldehyde + syngas). A number of chemicals, including acetic acid, methyl acetate, acetic anhydride and vinyl acetate, are of great commercial importance and will be discussed here. They all contain the acetyl group (CH,CO, often designated as AC). 7.3.1. Formaldehyde Formaldehye (HCHO) production, worldwide, has historically been the largest single application of methanol. Methanol constitutes 35.3 wt% of MTBE and its use in the manufacture of this and other methyl ethers as octane enhancers and fuels became, in 1944, the chief outlet for methanol. Formaldehyde is used in a large variety of resins including urea-formaldehyde polymers, in addition to its use in the manufacture of pentaerythritol, hexamethylenetetramine and other chemicals (Calkins, 1984). Formaldehyde is manufactured commercially by air oxidation, mostly by oxidation of methanol over a silver catalyst at atmospheric pressure; a 99% conversion of methanol is achieved with a molybdate catalyst at about 400°C and one atmosphere. A selectivity of about 95% is achieved at a 99% conversion of methanol (Machiels et al., 1984). CH,OH + l/20, + HCHO + H,O (7.1)

I.

7.3.2.

Acetic

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acid

Acetic acid has been produced in large quantities for more than 100 years. At present, more than five million tons per year worldwide are manufactured (Sheldon, 1994). Its main uses are in the manufacture of cellulose acetate, vinyl acetate monomer, acetic anhydride, a number of acetates, pharmaceuticals, dyes, pesticides, the dimethyl ester of terephthalic acid, as a solvent in xylene oxidation, and in other applications. Changes in the various methods of manufacture of acetic acid reflect underlying trends in the chemical industry toward cheaper, more available feedstocks, more energy efficient processes and the discovery of new, more active catalysts. Excellent reviews of the synthesis of acetic acid are available (Whyman, 1985; Papp and Baems, 1991; Parshall and Ittel, 1992; Agreda and Zoeller, 1993). In some countries, small amounts of acetic acid are still obtained by fermentation of ethanol or by the distillation of wood. The first large synthetic process for manufacture of acetic acid (AcOH, where AC is CH,CO-) was based on the reaction of the expensive and energy-rich compound acetylene to acetaldehyde followed by oxidation (Eq. (7.2)) (Aquilo et al., 1983). HC = CH + H,OHz+CH,COOH There are now three major processes for the production oxidation via the Wacker process (Eq. (7.3)) (Jira et al., butane or naphtha to acetic acid in 30-60% yields (Parshall and the carbonylation of methanol to acetic acid (Eqs. (7.5)

(7.2) of acetic acid: (a) ethylene 1976); (b) the oxidation of and Ittel, 1992) (Eq. (7.4)); and (7.6)).

CH, = CH, + [0] ?CH,COOH

(7.3)

Butane or naphtha + [0] + mixed acids

(7.4)

CH,OH

+ CO$CH,COOH

(7.5)

CH,OH

+ COzCH,COOH

(7.6)

BASF discovered and commercialized the cobalt iodide-catalyzed carbonylation of methanol in the mid-1960s (Reppe et al., 1961; Hohenschutz et al., 1966). However, the cobalt-catalyzed carbonylation of methanol to acetic acid operated under severe conditions of temperature and pressure and it was quickly replaced by a rhodium iodide catalyst system (Paulik and Roth, 1968). This process, known as the Monsanto process, operated under milder conditions with greater selectivity than the cobalt based process. Rh is recovered in over 99% yield. A comparison of the rhodium and cobalt systems is given in Table 18. The Monsanto process (Roth et al., 1971; Paulik, 1973) can produce acetic acid at 1 atm, but the commercial synthesis is carried out at 30-40 atm and 180°C with Rh as catalyst and HI ( = 10-l M) as promoter. The catalytic system is very corrosive, requiring expensive steels for construction; complete recycle of Rh and HI is necessary. Iodine is used to convert methanol to the more electrophilic methyl iodide. A catalytic cycle for the rhodium-catalyzed carbonylation of methanol is shown in Fig. 34 and a

214 Table 18 Comparison

1. Wender / Fuel Processing

of rhodium-

and cobalt-catalyzed

Metal concentration Temperature Pressure Selectivity (on CH,OH) By-products

Technology 48 (1996) 189-297

syntheses

of acetic acid

Rhodium

Cobalt

= IO-’ M = 180°C 30-40 atm 0.99 0

= lo- ’M = 220°C 700 atm 0.9 CH,, CH,CHO, C,H,OH, CO, C, H=,CO,H, acetates, 2-ethylbutanol

conceptual flowsheet for the carbonylation is given in Fig. 35. The reaction is zero order with respect to methanol and carbon monoxide and first order in methyl iodide (Roth et al., 1971). The active species has been identified as cis-Rh(CO),I; by infrared analysis. Large excesses of water are used, supposedly to suppress formation of methyl acetate. It has been found that the addition of lithium iodide allows for a great reduction in the amount of water required while maintaining high rates (Zoeller, 1993). At present, about half of acetic acid production worldwide uses the Monsanto rhodium-iodide catalyst. This route to acetic acid has very favorable economics and essentially all new plants will use the Rh-catalyzed methanol carbonylation. Indeed, carbonylation of methanol to acetic acid is one of the most successful chemical processes that is based solely on the use of syngas. At present, 59% of worldwide acetic acid output is made by this process. In 1986, Monsanto sold the rights to the methanol carbonylation process for acetic acid production to BP Chemicals (Howard et al., 1993). In 1996, BP introduced a new catalyst system based on the use of iridium acetate enhanced with various promoters. It is claimed that the iridium-based route is more efficient, cutting energy and purification costs compared with the rhodium-catalyzed methanol carbonylation (Chem. Eng. News, 1996). It is of interest that earlier work showed that iridium is an active catalyst for methanol carbonylation but is not as selective as rhodium. There have been attempts to use a cheaper nickel based catalytic system for the carbonylation of methanol (Rizkalla, 1987). This is still a possible catalytic route but the hazards of nickel tetracarbonyl and the complex nature of the nickel catalyst, in addition

H.0

Fig. 34. Mechanism

for the rhodium-catalyzed

methanol carbonylation

(Zoeller,

1993).

I. Wender/Fuel

Procrssing

27s

Technology 48 (1996) 189-297

-

Fig. 35. Conceptual

flowsheet for the carbonylation

of methanol to acetic acid (Zoeller,

1993).

to the incidence of side reactions such as hydrocarbon formation, will probably delay or possibly prevent commercialization of a nickel based methanol carbonylation process. Methanol can be converted to acetic acid via methyl formate in a two-step process shown by the following reactions (Lee et al., 1990). base

CH,OH+CO-+HCOOCH,

U-7)

HCOOCH,

(T-8)

+ CH,COOH

Many homogeneous and heterogeneous catalysts have been used for this conversion to acetic acid (Torrence et al., 1991); all involve addition of methyl iodide. Most of the chemistry is similar to that already proposed for the direct carbonylation of methanol and is related to the commercial acetic acid process. It is not likely that this route to acetic acid through methyl formate will be used industrially; it converts methanol to acetic acid in two somewhat difficult steps. 7.3.3. Acetic anhydride About a million tons of acetic anhydride are produced each year (Cook, 1993). It is used to acetylate the hydroxyl group of a number of substrates to yield products such as cellulose acetate (its chief use>, textile fibers, cigarette filters, various polymers, coatings, aspirin, acetaminophen, flavors, fragrances, sweeteners and other uses including as a general reagent for making acetates as a replacement for acetyl chloride (which releases the corrosive HCl).

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The old process for production of Ac,O used the energy intensive ketene route shown in the following equations for over 65 years to manufacture acetic anhydride. The ketene is made from either acetone (Fife and Zhang, 1986) or acetic acid (Cook, 1993). CH,COCH, CHJOOH

750°C

+ CH, = C = 0 + CH,

(7.9)

Et.&‘%

+

CH, = C = 0 + H,O

(7.10)

700°C

CH, = C = 0 + CH,COOH + (CH,CO),O

(7.11)

The latest and seemingly most viable synthesis uses a homogeneous catalysis system to convert methyl acetate to Ac,O CH,COOCH,

+ CO cat~st(CH3CO),0

(7.12)

Here a carbonyl group is apparently inserted between the 0-CH, bond of methyl acetate. The mechanism for this insertion is complex, involving a noble metal, methyl halide/ionic halide and an acetate salt. This synthesis uses the Tennessee-Halcon process to convert methyl acetate to acetic anhydride. In 1983, the Eastman Chemical Company in Kingsport, Tennessee became the first major chemical company to begin again to use coal as a raw material for the manufacture of industrial chemicals on a large scale (Agreda et al., 1992). They began manufacturing AC*O from bituminous coal. The Eastman Chemical Company uses two Texaco gasifiers to produce syngas from coal. The process is based on technology developed independently by Halcon SD and by Eastman in the 1970s (Worsham, 1994). As shown by the following equations, the entire synthesis of Ac,O is developed from a syngas feed. CH,OH + CO --) CH,COOH CH,OH + CH,COOH + CHJOOCH, CH,COOCH,

(7.13) + H,O

+ CO --) (CH,CO),O

(7.14) (7.15)

There are important differences between the aqueous acetic acid synthesis and the anhydrous Ac,O system (Zoeller et al., 1992; Worsham, 1994). It is necessary to have an iodide salt and a reducing agent (H,) present for the commercial synthesis of Ac,O. A significant process implication is that the carbonylation of methyl acetate has a lower heat of reaction; this lower thermodynamic driving force results in equilibrium limitations for Ac,O conversion. The halolysis of methyl acetate by LiI is the key for the increase in activity brought about by addition of inorganic promoters (Luft and Schrod, 1983). CH,COOCH,

+ LiI -+ CH,I + LiOOCCH,

(7.16)

Eastman will produce about 1.2 billion pounds per year of Ac,O from 1100 tons of bituminous coal per day, thus becoming the world’s largest manufacturer of acetyl products and eliminating the equivalent use of about 1.23 million barrels of oil per year.

1. Wrnder / Fuel Processing

To sum up, the Eastman sequence.

Technology 48 (1996) 189-297

process proceeds according

to the following

277

overall reaction

Cu/ZnO

2H, + CO

+

CH,OH

(7.17)

Rh.I-

CH,OH

+ CO

+ I6OT.25

CH,OH

+ CH,COOH

CH,COOCH, (CH,CO),O

CH,COOH

(7.18)

am

(7.19)

--) CH,COOCH,

+ CO Rhc~p’rx(CH~CO),O + cellulose

(7.20)

-+ cellulose acetate + CH,COOH

(7.21)

(recycled)

Eastman has undergone a recent expansion and is well on its way to establishing economic routes to the production of a series of industrial oxygenated chemicals with this background. An overall flow diagram for Eastman’s acetic anhydride complex is shown in Fig. 36. There are upcoming developments in this area of chemicals production (Papp and Baems, 1991). Hoechst has patented a process to co-produce acetic acid and Ac,O by carbonylation of acetic acid and methyl acetate in a single reactor. Rhodium salts plus ammonium or phosphonium iodide make up the catalyst system. BP has announced a similar process. The different processes are compared in Table 19. 7.3.4. Vinyl acetate The largest single use of acetic acid, almost six billion pounds per year, is in the production of vinyl acetate (Summer and Zoeller, 1993). Vinyl acetate-derived polymers are found everywhere in our modem society. Vinyl acetate gives us polyvinyl acetate,

COal Gasification Plant

suthr

Acenc Acid

Fig. 36. Block diagram of main steps required to produce acetic anhydride

from coal (Cook, 1993).

I.

278

Table 19 Reaction conditions

Temperature Pressure Catalyst

Wemler / Fuel Processing Technology 48 (1996) 189-297

for the carbonylation

of methanol

and methyl acetate

Acetic acid Monsanto

Acetic anhydride Eastman

Acetic acid, anhydride Hoechst

Acetic acid, anhydride BP

190- 195°C 30-35 bar Rh

190°C

> 150°C

50 bar Rh

> IO bar Rh

183°C 30 bar Rh

used in adhesives, coatings, latex paints and finishes; polyvinyl alcohol, used in textile sizing, paper-coating, pastes, plywood adhesives and cement additives; ethylene-vinyl acetate polymer for packaging film, hot-melt adhesives and cable coverings; ethyl-vinyl alcohol copolymer for flexible containers and bottles; and polyvinylbutyral for laminated safety glass, automotive and architectural applications. Polyvinyl acetate formed for magnet/wire insulation and vinyl chloride-vinyl acetate copolymer for protective surface coatings are important uses of vinyl acetate in our surroundings. Since 1960, most vinyl acetate is produced by the palladium-catalyzed reaction of ethylene and acetic acid, involving the subsequent reoxidation of the reduced palladium (Robinson, 1965; Daniels, 1978). CH, = CH, + CH,COOH + l/20,

ZCH,

= CHOCOCH, + H,O

(7.22)

vinylacetate

As acetic acid is made from methanol and CO, this process is based on syngas for 70% of its weight, giving overall yields of 90% and 95% based on ethylene and acetic acid respectively (Whyman, 198.5). Halcon has obtained patents describing a syngas based process leading to vinyl acetate (Parshall and Ittel, 1992): methyl acetate (or DME) reacts with syngas to form CH,CH(OAc), (EDA), which loses acetic acid to give vinyl acetate. H,/co - AcOH CHJOOCH, --, EDA + CH, =CHOCOCH, (7.23) metal oxide

Eastman and Halcon have piloted routes from acetic anhydride to vinyl acetate but the processes are not economical at present. These syntheses of vinyl acetate based on syngas chemistries are still under development. A process operated by Celanese reacts acetaldehyde and acetic anhydride (Ac,O) to form ethylidene diacetate, which is then cracked to vinyl acetate over a sulfonic acid catalyst (Rizkalla and Goliaszewski, 1987) CH,CHO + Ac,O + (AcO),CHCH,

(7.24)

(Ac,O),CHCH,zAcOCH

(7.25)

heat

= CH, + CH,COOH vinyl

acetate

7.3.5. Homologation (reductive carbonylation) of methanol to ethanol Wender et al. (1949, 1951) showed that Co,(CO), catalyzed the reaction of methanol with syngas at 185°C and 270 atm to yield ethanol in 39% selectivity and 70% methanol conversion. CH,OH + 2H, + CO --, CH,CH,OH + H,O (7.26)

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The reaction is really a hydrocarbonylation of methanol but the term “homologation” has come to mean the reaction of an organic molecule with syngas resulting in the extension of the carbon chain by one methylene (-CH,-1 unit. The homologation of methanol to ethanol received little attention following its discovery. In the 197Os, interest was rekindled as a result of the oil crises. Numerous reviews of further work in this area have since appeared (Albanesi, 1973; Wender, 1976; Piacenti and Bianchi, 1977; Koermer and Slinkard, 1978; Pretzer and Kobylinski, 1980; Chen et al., 1982; Fakley and Head, 1983; Jenner, 1989; Parshall and Ittel, 1992). The homologation of methanol is thermodynamically favored as the pressure is raised, but the chief problem is kinetic control over the number of reaction products, including methane, acetates, acetic acid, acetals and CO,. Cobalt complexes, first used, are good catalysts in their activity and selectivity for the conversion of methanol and syngas to ethanol and acetaldehyde. Rhodium, usually more active than cobalt in homogeneously catalyzed reactions involving methanol and syngas. produces acids and esters, with ethanol a significant product at high partial pressures of hydrogen (Dumas et al., 1979). Keim (1989) has been studying the conversion of methanol to acetaldehyde H,,‘CO

CH,OH

-+

CH,CHO

(7.27)

In the presence of a polar solvent (dioxane, sulfolane), a catalyst system [(Ph,P),N] [Co(CO),]/12 gave a methanol conversion of 97% with 80% selectivity to the aldehyde. A pilot plant was operated for over three years in cooperation with Union Rheinische Braunkohle AG Wesseling. Keim feels that an acetaldehyde plant based on the homologation of methanol may be built in the near future. An engineering evaluation concluded that this route to the aldehyde is more economical than the Wacker process (Jira et al., 1976). However, additional capacity for the aldehyde is not presently needed because its principal use, oxidation to acetic acid, has been largely replaced by the Monsanto methanol-to-acetic acid process. A strong increase in activity to ethanol has been obtained by adding iodide promoters of methanol to such as I, or CH,I (Gauthier-Lafaye et al., 1982). The homologation ethanol using phosphine ligands and Ru compounds has been described by Fiato (1980) and by Bozik et al. (1980). With a catalyst containing Co, Ru, I and a phosphine ligand, methanol conversion was 50-60% (Papp and Baems, 1991). The relative rates of homologation of several alcohols with cobalt carbonyl and iodine was studied by Berty et al. (1956). Ethanol reacts 42 times more slowly than methanol (Jenner, 19891, consistent with an S, 2 displacement of iodide by Co(CO),. Propanol and isopropanol are even less active. However, t-butyl alcohol, which readily forms carbenium ions, reacts very rapidly, evidently by an S,l mechanism, to form (CH,),CCH,OH. There is no commercial plant that now uses the reductive carbonylation of methanol to obtain either acetaldehyde or ethanol. Poor selectivity and severe operating conditions are obstacles to the use of this route to ethanol. Although the conversion of starch to ethanol by fermentation has taken over, there are those who feel that the catalyzed homologation of methanol to ethanol has a future.

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Processing Technology 48 (1996) 189-297

7.3.6. Homologation of ethers, esters and carboxylic acids

Braca et al. (1978), Braca et al. (1981) found that dimethyl ether and methyl acetate, which are formed in the homologation of methanol, could be carbonylated further CH,OCH, + 2H, + 2C0 CHJOOCH,

200°C

--, CH,COOC,H,

+ H,O

(7.28)

+ 2H, + CO --) CH,COOC,H,

+ H,O

(7.29)

I50 atm

These reactions were achieved with Ru/Icatalysts. Ethyl acetate selectivities reached 80%. Cobalt and rhodium did not carbonylate these substrates (Sheldon, 1983). The same type of catalyst, Ru/I-, is used in the conversion of carboxylic acids. Acetic acid, for instance, yields propionic acid plus small amounts of n-butyric and pentanoic acids (Knifton, 1981a, Knifton, 1981b, Knifton, 1981~). CH,COOH

Hz/co Hz/co --$ CH,CH,COOH + CH,CH,CH,COOH,etc.

(7.30)

RU/I_

7.3.7. Methyl formate Methyl formate is an intriguing chemical that has been proposed as a building block in C, chemistry (Keim, 1983, Keim, 1987). The paper by Lee et al. (1990) provides an excellent review of the synthesis, properties, and uses of methyl formate. At present, however, methyl formate does not qualify as a major building block in C , chemistry, but an examination of the various methods of its synthesis and of its possible uses as a feedstock for conversions to a number of important chemicals warrants discussion. Methyl formate is convenient to handle, store, and transport as a stable liquid; it boils at 31.5”C and is easily separated from methanol, b.p. 64.7”C. It could be handled, stored and transported in a manner resembling that for liquefied petroleum gas (LPG). Methyl formate is synthesized commercially by the alkali methoxide-catalyzed reaction of methanol and CO (Christiansen, 1919). Methoxide ion (OCH,)- is a strong base and nucleophile and it attacks the electrophilic CO molecule CH,O- + C = 0- + (CH,OCO) (CH,OCO)-

+ CH,OH + HCOOCH, + CH,O-

(7.31) (7.32)

NaOCH,

CH ,OH + CO

+

HCOOCH,

(7.33)

A drawback, however, is that dry methanol must be used in this synthesis. The CO stream containing at least 50% CO and not more than 10 ppm of water or 50- 100 ppm of CO, should be available at about 600 psig. Conversion of methanol and CO is typically 30% and 95%, respectively; selectivity to methyl formate is = 99%. The presence of hydrogen or nitrogen in the feed gas does not interfere. Methyl formate itself has limited application as a solvent and as an insect control agent. It is mainly used as an intermediate in the production of formic acid and formamides. There are a number of alternate ways of synthesizing methyl formate that are of interest. The Mitsubishi Gas Chemicals Co. had a unit that produced over 20000 tons

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281

per year (tpy) of methyl formate by the dehydrodimerization of methanol (Komatsu and Yoneoka, 1988; Nojiri and Misono, 1993). The plant may have been abandoned because of the rising price of methanol. 2CH,OH

+

HCOOCH,

+ 2H,

(7.34)

250 - 300°C 3-5atm Cu/Zr/Al,O,

Perhaps the most promising route to methyl formate is by its direct synthesis from syngas by the Brookhaven and related processes or by the concurrent synthesis of methanol and methyl formate. In each case, methyl formate can be the principal product, separated from by-product methanol by distillation. 2H, + 2C0 -+ HCOOCH,

(7.35)

Good yields of methyl formate and methanol can be achieved at 50-100°C and = 600 psig. Significant improvements in these processes may be achieved in the future. Methyl formate has a number of potential industrial uses. It is sometimes necessary to separate H, and CO in syngas or to adjust their ratio. Under certain conditions CO and/or H? must be recovered from non-conventional sources such as off-gases from the production of steel or from vent streams from chemical processes. The commercial carbonylation route to methyl formate can be operated with as little as 50% CO in the feed gas, as mentioned earlier. The syngas can be enriched in CO in a once-through methanol synthesis unit or by removal of H, by a membrane, by diffusion or by some other means (Halcon SD, 1985). The CO-rich syngas is carbonylated to methyl formate, which could be stored for further utilization including decomposition to pure CO. The hydrogen-rich vent stream has many uses (Lee et al., 1990). Methyl formate can be decomposed to give methanol and CO HCOOCH,

-+ CH,OH

+ CO

(7.36)

Under these conditions methyl formate decomposes selectively even in the presence of excess methanol. With a KCl/activated carbon catalyst, 99.5% of methyl formate is converted to methanol and CO with a selectivity exceeding 99%. Methyl formate is thus a good CO carrier.A 1: 1 syngas can be obtained from methyl formate by treating it at 300-350°C over alkali or alkaline earth oxides HCOOCH,

--+ 2H, + 2C0

(7.37)

This syngas has the stoichiometry and purity that makes it a suitable feed for chemical syntheses based on a temporarily out-of-commission coal gasifier or for a hydroformylation unit. Methyl formate has been reported to isomerize selectively to acetic acid (Pruett and Kacmarik, 1982). HCOOCH,TCH,COOH

(7.38)

No CO is consumed in this reaction but it only takes place under CO pressure. Although many transition metals catalyze this isomerization, it appears that a Rh-LiI

282

1. Wender/Fuel

Processing Technology 48 (19961 189-297

catalyst is highly efficient for the reaction at 180°C. The methyl formate conversion and the molar selectivity to acetic acid are both over 99%. Lithium iodide aids the reaction by cleaving methyl formate and the intermediate methyl acetate to CH,I (Schreck et al., 1988). Methyl iodide is an intermediate both in the direct synthesis of acetic acid from methanol and CO and in the isomerization of methyl formate to acetic acid. The difference is that the methyl formate synthesis involves insertion of CO into an oxygen-hydrogen bond whereas acetic acid is formed by insertion of CO between the carbon-oxygen bond in methanol. This isomerization can find use when methyl formate is an undesirable by-product (as in the oxidation of butane) or if the formate could be produced more cheaply. In some locations, dehydrogenation of methanol to methyl formate, with subsequent isomerization, could be an economic route to acetic acid. It is unlikely, however, that this isomerization, which has been known for many decades (Dreyfus, 1929; Roper et al., 1985), can compete with the Monsanto process for the synthesis of acetic acid from methanol and carbon monoxide. 7.3.8. Formic acid Until recently, most US formic acid production was a by-product from the liquid-phase oxidation of butane to acetic acid. With the advent of cheaper acetic acid made by carbonylation of methanol, the importance of this by-product route from hydrocarbons has declined. The present method for the synthesis of formic acid is by hydrolysis of methyl formate but, as retroesterification occurs readily, HCOOH is generally made by first synthesizing formamide (HCONH,) (Peltzman, 1984) HCOOCH, + NH,

80- IOO”C

+

HCONH, + CH,OH

5 atm

(7.39)

The formamide is then continuously hydrolyzed to HCOOH HCONH, + H,O + HCOOH + NH,

(7-40)

About 400000 tons per year of formic acid are produced worldwide. Its major industrial uses are in the textile and leather industries. It has been used in Europe for many years as a silage preservative. 7.3.9. Ethylene glycol The direct synthesis of ethylene glycol from syngas is a most attractive route from a raw material point of view (Table 171, but this has been an elusive target. At present, ethylene glycol is manufactured commercially by the reaction of ethylene with oxygen over a silver catalyst. The ethylene oxide so formed is then hydrolyzed to ethylene glycol. Over five billion pounds of ethylene glycol were produced in the United States in 1993. The direct production of ethylene glycol from syngas was discovered by DuPont using cobalt as the catalyst (Aquild et al., 1983). More recently, Union Carbide synthesized ethylene glycol directly from syngas with a rhodium catalyst, but the reaction rate is too low, requiring severe conditions and difficult separation of the glycol

I?83

1. Wender / Furl Proc’rssin~ Technology 48 (I 996) 189-297

from co-produced methanol, propylene process has been carried to a semi-works

glycol, glycerol and other by-products. but is not economic at present.

The

Kh

3H,+CO

-+

(7.41)

HOCH,CH,OH

210- 250°C 500 - 3400 atm

Many companies are endeavoring to find a new commercial process for the synthesis of ethylene glycol. As shown in Fig. 33, oxidative carbonylation of methanol to dimethyl oxalate (itself useful in agriculture as a compound which releases nitrogen slowly, in the food industry and in pharmaceuticals) followed by hydrogenation yields ethylene glycol. Perhaps the most promising process involves the synthesis of dimethyl oxalate by oxidation of methanol with oxygen and nitric acid, a reaction applied commercially since 1978 by Ube Industries/Japan (Roper, 1991; Forster, 1976; Wallet-, 1985). 2CH,OH

+ 2N0 + l/20,

-+ 2CH,ONO

(7.42)

+ H,O

Anhydrous methyl nitrite is carbonylated in a second stage over a Pd/Fe furnishing a 97% yield of dimethyl oxalate, which is then hydrogenated ruthenium catalyst to ethylene glycol in 90% yield. 2CH,ONO CH,OOC

+ 2C0 + H,COOC - COOCH,

+ l/20,

+ 2N0

+ 4H, + HOCH,CH,OH

The nitric oxide (NO) is converted NO + CH,OH

- COOCH,

+ 2CH,OH

catalyst, over a (7.43) (7.44)

to methyl nitrite for recycle

--) CH,ONO

+ 1/2H,O

(7.45)

This is an attractive route to ethylene glycol involving methanol, CO and HZ in different stages. Formaldehyde derived from methanol can be converted to glycolic acid derivatives (Fig. 33) which can be hydrogenated to ethylene glycol. Other routes to the glycol by reaction of formaldehyde have been studied by Celanese (Kollar, 1982), by Monsanto and by Exxon (Chem. Eng. News, 1983). 7.3. IO. Dimethyl carbonate Dimethyl carbonate is a non-toxic chemical that is being promoted as a replacement for the toxic intermediates phosgene and dimethyl sulfate. This follows the current trend in the chemical industry to reduce the use of highly toxic substances. A glance at Fig. 33 indicates that dimethyl carbonate can be synthesized by the oxidative carbonylation of methanol. Roman0 et al. (1980) synthesized dimethyl carbonate from methanol, carbon monoxide and oxygen in the liquid phase in a slurry reactor using a cuprous chloride catalyst 2CH,OH

+ CO -I- l/20,

-+ (CH,O),CO

+ HZ0

(7.46)

EniChem has produced dimethyl carbonate in a 5000 tons per year plant since 1983 (Mauri et al., 1985). The plant output was expanded to 8800 tons per year in 1988 to meet the growing interest in the carbonate. The chief captive uses are as a phosgene substitute in monomers for optical resins and in the manufacture of resins for polycar-

284

I. Wender/

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bonates and carbamates; it is also used in the synthesis of synthetic lubricants and in polyurethane manufacture. Dimethyl carbonate is used in place of the toxic dimethyl sulfate in methylation reactions. It is being considered in fuel reformulation as an oxygen-rich octane enhancer in motor fuels. Dow has also developed a process for the synthesis of dimethyl carbonate (Chem. Eng. News, 1987). The carbonate is produced via a vapor-phase carbonylation of methanol using a solid catalyst of activated carbon impregnated with cupric chloride. It is claimed that this process reduces corrosion problems and simplifies product recovery. 7.3.11. Methyl methacrylate Methyl methacrylate polymers have many end uses and about one million tons per year are manufactured worldwide. The o-methyl group of polymethylacrylates imparts the stability, hardness and stiffness of methacrylate polymers. The current process uses an acetone cyanhydrin route which utilizes methanol in its manufacture (Porcelli and Juran, 1986). Environmental pressures have resulted in on-site synthesis of HCN to avoid shipment of by-product HCN (from manufacture of acrylonitrile). To avoid the use of HCN, methyl methacrylate can be manufactured from propylene C,H, + l/20,

+ CO + CH,OH 4 CH, = C(CH,)COOCH,

+ H,O

(7.47)

or C,H, + 0, -I-CO + CH,OH --, MMA + 2H,O

(7.48)

Carbon monoxide is expensive to transport and the amount needed for a reasonably sized methyl methacrylate plant is much less than the economic size for low cost CO production. However, a propylene carbonylation route to MMA would make sense as part of a complex that produced other C , derivatives. One plant in Germany manufactures methyl methacrylate from propionaldehyde, which is produced from ethylene via the hydroformylation of ethylene C,H, + CO + HCHO + CH,OH + MMA + H,O

(7.49)

A disclosure from SRI International reports that Shell will produce MMA from propyne (CH, = CH) at a cost lower than that of the acetone cyanohydrin and other processes (Chemical Marketing Reporter, 1993). Although Shell has not yet announced any such plans, it is indicated that the process is described in US and European patents, has been piloted and is ready to go into commercial production. Evidently, the MMA would be produced according to the following equation CH,C = CH + CO + CH,OH --) CH, = C(CH,)COOCH, The propyne (methylacetylene) ethylene plant.

(7.50)

is obtained from the C, product stream from an

7.3.12. Chlorinated hydrocarbons Methyl chloride, methylene chloride, chloroform and carbon tetrachloride are derived mainly by reaction of methanol and chlorine. The two principal processes for the commercial production of methyl chloride are the chlorination of methane and the

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reaction of methanol and hydrogen chloride. With methanol, the process produces and recycles the HCl. In the chlorination of methane, I mol of HCl has to be disposed of, generally as 3 1 wt% NC1 (Calkins, 1984). In the US, the liquid-phase reaction of methanol and HCl has the widest use. In Europe and Japan, the gas-phase methanol hydrochlorination process is used more than in the US (Holbrook, 1993). With methanol as starting material, the reactions that occur are CH,OH

+ HCl -+ CH,CI

+ H,O

(7.51)

CH,Cl

+ Cl, -+ CH,Cl,

+ HCI

(7.52)

CH,CI

+ 2C1, + CHCl,

+ 2HCl

(7.53)

CH,Cl

+ 3C1, --) Ccl,

(7.54)

+ 3HCI

7.3.13. Me~hylamines The methylamines have a number of applications, with dimethylamine used the most, monomethylamine the next and trimethylamine the least. They are produced by the reaction of methanol and ammonia in the vapor phase using a dehydration catalyst at about 450°C and l-20 atm (Calkins, 1984). CH,OH

+ NH, + CH,NH,

2CH,OH

+ NH, -+ (CH,),NH

3CH,OH

+ NH, -+ (CH,),N

+ H,O + 2H,O + 3H,O

(7.55) (7.56) (7.57)

Depending on the amine product desired, a 2: 1 to 6: 1 ratio of ammonia to methanol is employed. The reaction reaches approximate thermodynamic equilibrium as a function of the ammonia to methanol ratio. The more utilized dimethylamine is produced in only 30-40% of the total amine yield. The other amines, plus additional ammonia, are recycled to the reactor. Yields based on methanol can reach 99%. Silver/alumina catalysts with silver phosphate, molybdenum sulfide or cobalt sulfide promoters are the catalysts of choice.

8. Miscellaneous reactions of syngas or carbon monoxidekarbonylation) 8.1. Introduction The major reactions of syngas and of CO (carbonylation) have been discussed but there are other reactions of these gases which are too numerous to detail. Some of these are used commercially today, hundreds of others appear interesting but have not been exploited (Colquhoun et al., 1991). A number of these reactions were used in industrial processes that have been improved or replaced by cheaper or environmentally superior syntheses. Many of these reactions of syngas or of CO are proprietary and it is difficult to obtain detailed information on their commercial viability. There is sometimes a fine line between commercial and pre-commercial processes.

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8.2. Reppe carbonylation reactions

Reppe and his coworkers, between 1938 and 1945, invented a series of reactions of acetylenes and olefins with CO and protic reactants such as water, alcohols or amines to give carboxylic acids and their derivatives (Reppe, 1953; Reppe and Vetter, 1953; Falbe, 1970). The applicability of these syntheses was amazingly general. At one time, some 600000 tons of Reppe carbonylation products were made per year (Mullen, 1980). The following are examples. HC = CH + CO + H,O --) CH 2 = CHCOOH (acrylic acid)

(8.1)

HC = CH + CO + ROH + CH, = CHCOOR

(8.2)

HC = CH + CO + R,NH + CH, = CHCONR,

(8.3)

The same types of reactions could be carried out with olefins CH, = CH, + CO + ROH + CH,CH,COOR

(8.4)

CH, = CH, + CO + R,NH + CH,CH,CONH,,etc.

(8.5)

The carbonylation of acetylenes and olefins could be carried out stoichiometrically with metal carbonyls at atmospheric pressure with the carbonyl acting as both a supplier of CO and a catalyst. Commercially, however, the syntheses were carried out at high CO partial pressures with catalytic amounts of transition metal carbonyls produced in situ from various metal salts; nickel compounds were most effective. Acetylenes react more easily than olefins, which require higher temperatures and partial pressures than acetylenic feedstocks. Newer plants for the synthesis of acrylic acid use a process based on the oxidation of propylene H,C = CH - CH, + 0, tc? H,C = CHCHO -H,O

(8.6)

The aldehyde is then oxidized in air with a molybdenum catalyst to acrylic acid, CH, =CHCOOH (Weissermel and Arpe, 1993). The syntheses of acrylic acid and acrylate esters are examples of Reppe reactions. These compounds are used in the manufacture of polymers and are made from acetylene, CO and water or an alcohol. Acrylic acid is formed in a catalytic process by treating acetylene in tetrahydrofuran with CO, water and NiBr, at about 200°C and 60 atm (Toepel, 1964). Although this process has been displaced by one based on the oxidation of propylene in the US, it is still used elsewhere to a considerable extent (Weissermel and Arpe, 1993). Reppe and coworkers accumulated enough experience to be able to carry out reactions of acetylene itself under elevated pressures without incident. There have been several instances in which explosions have occurred when acetylene reactions were carried out at elevated pressures and temperatures. Worldwide capacity for acrylic acid is about two million tons per year.

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287

acids

In the Koch reaction, olefins are converted to highly branched carboxylic acids using strong acids such as H,SO, or H,PO,/BF, as catalysts. The olefin adds a proton to form a carbenium ion, which is then isomerized + RdHCH,CH

RCH,CH = CH,H-;RCH,;HCH, RC(Cw,),

j

-+ &(CH,)*

(8.7)

+ CO + H,O 2o-Jo”c RC( CH,),COOH 20- lOOatm

Esters of these branched so-called neo-acids exhibit unusual thermal and oxidative stability and are difficult to saponify. These properties are largely the result of alkyl branching LYto the carboxyl group. Based on their great stability, the acids are suitable components of synthetic oils. Several companies now operate industrial processes based on this reaction. 8.4. Homologation

of carboxylic

acids

Knifton (1981a), Knifton (1981b) has discovered a useful way of lengthening the carbon chain of carboxylic acids by adding one carbon at a time. The reaction, using acetic acid as an example, follows RuO,, HI

CH,COOH+CO+2H,

CH,CH,COOH

+

+ H,O

1OOatm.22OT

(8.8)

With propionic acid, two isomeric products are obtained CH,CH,COOH

+ CO + 2H, + CH,CH,CH,COOH

+ (CH,),CHCOOH (8.9)

The normal to iso acid ratio varies from 4 to 8. Since acetic acid is itself synthesized entirely from syngas, all these acids can be built exclusively from syngas with no need for petrochemical feedstocks. 8.5. Carbonylation

of nitro compounds

to isocyanates

Nitroaromatic compounds react with carbon monoxide to give the corresponding isocyanates in the presence of palladium catalysts (Parshall and Ittel, 1992). With nitrobenzene, the reaction may be written C,H,NO,

+ 3C0 2 C,H,N = C = 0 + 2C0,

(8.10)

190°C

This is a reductive carbonylation reaction and it has advantages over the usually practiced two-step syntheses of isocyanates from the toxic compound phosgene (COCI, > C,H,NO,

%6H$JH~C~‘2C6H5NC0

+ 2HCI

(8.11)

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The reaction given in Eq. (8.10) is an environmentally better way to prepare diisocyanates such as toluene diisocyanate (TDI) and methylene diphenylisocyanate (MDI), which are used to manufacture polyurethane for fibers, coatings and adhesives; their largest use is in the preparation of foams. The structures of MD1 and TDI are given below (Scheme 1) Most modem polyurethanes are prepared from diisocyanates such as TDI or MD1 in combination with diols such as ethylene or propylene glycol, polyether diols, aliphatic polyester diols or siliconediols. 8.6. Amidocarbonylation

reactions

In general, many carbonylation reactions are being used for the production of high-value specialty chemicals. As an example, the synthesis of amido acid precursors from olefins or aldehydes plus syngas and amides has been under study by Texaco (Lin and Knifton, 1992; Knifton et al., 1993). Wakamatsu et al. (1971) discovered the amidocarbonylation reaction but this potentially useful reaction was not thoroughly investigated until recently. Knifton and coworkers have studied the chemistry involved to synthesize a variety of specialty chemicals by tailoring cobalt and rhodium catalysts to make particular products selectively. These amidocarbonylation products can be formed from a wide variety of olefins and aldehydes which react with syngas to give useful products, including specialty surfactants, surface active agents (C ,4-C ,6 alkyl amido acids), food additives, chelating agents, and intermediates for sweeteners such as aspartame. With available inexpensive C ,2-C 14 straight-chain terminal olefin feedstocks, amido acids with up to 95% linearity can be synthesized Co,Rh, 100°C

CH,(CH,),,CH CH,(CH,)

= CH, + CH,CONH,

+ H, + CO

,,CH(COOH)NHCOCH,

+ 50- 130atm (8.12)

8.7. Esters from epoxides An example of an interesting carbonylation reaction which occurs from easily available compounds under comparatively mild conditions (6X, 130 atm) is the

Scheme I.

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synthesis of B-hydroxyesters from epoxides. Carbonylation of ethylene oxide or propylene oxides in methanol to the ester proceeds as follows (Heck, 1963; Eisenmann et al., 19611

RHF-,C”2

CofCO), -!- CO + CH,OH +

OH I RCHCH,COOCH,

(8.131

0

This synthesis

does not seem to have been generally

exploited.

8.8. Esters from ethers Ethers may be carbonylated to esters with both homogeneous transition metal catalysts (Piacenti and Bianchi, 1962). ROR + CO --) RCOOR

and heterogeneous (8.14)

Catalysts that have been used include Ni metal, Co/SiO, or NiI,/SiO,. Hydrogen halides such as HI promote the reaction, as in the conversion of methanol to acetic acid; the ether is probably cleaved by the acid promoter to yield alkyl halides and an alcohol. The alkyl halide is catalytically carbonylated to give an acyl halide that reacts with the alcohol to give the ester. This reaction may be applied to the conversion of dimethyl ether to esters ROR+HI*RRI+ROH

(8.15)

RI + CO + RCOI

(8.16)

RCOI + ROH + RCOOR + HX

(8.17)

Acknowledgements The support of the Electric Power Research Institute (EPRI) is gratefully acknowledged. Thanks are due to G.A. Mills for invaluable help and to J. Inga in preparation of the manuscript.

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