GB1566638A - Conversion fo coal to high octane gasoline - Google Patents

Description

(54) CONVERSION OF COAL TO HIGH OCTANE GASOLINE (71) We, SHELL INTERNATIONALE RESEARCH MAATSCHAPPIJ B.V., a company organized under the laws of the Netherlands, of 30 Carel van Bylandtlaan, The Hague, the Netherlands, do hereby declare the invention, for which we pray that a patent may be granted to us, and the method by which it is to be performed, to be particularly described in and by the following statement:- The invention relates to a process for preparing liquid hydrocarbons from coal.

Hydrocarbon mixtures boiling in the gasoline range can be obtained, for instance, by straight-run distillation of crude mineral oil, by conversion of heavy mineral oil fractions, for instance, by catalytic cracking, thermal cracking and hydrocracking and by conversion of light mineral oil fractions, for instance by alkylation.

In view of the increasing need of gasoline and the decreasing reserves of mineral oil there is a great interest in processes having the potentialities of converting carboncontaining materials not based on mineral oil, such as coal, in an economically justified way into hydrocarbon mixtures boiling in the gasoline range.

It is known that carbon-containing materials, such as coal, can be converted into mixtures of carbon monoxide and hydrogen by gasification. It is further known that mixtures of carbon monoxide and hydrogen can be converted into mixtures of hydrocarbons by contacting the gas mixtures with suitable catalysts. Finally, it is known that mixtures of paraffins and olefins boiling below the gasoline range can be converted into hydrocarbon mixtures boiling in the gasoline range by contacting the mixtures first mentioned with an alkylation catalyst.

The Applicants have carried out an investigation to examine to what extent the three above-mentioned processes can be used for preparing gasoline from coal. This investigation has shown that gasoline having a high octane number can be prepared from coal by combining the three abovementioned processes, provided that the following conditions are satisfied.

First of all, the gasification of the coal should be carried out at a temperature of from 1050 to 20000 C. From the mixture of carbon monoxide and hydrogen thus obtained an aromatic hydrocarbon mixture should then be prepared using a catalyst which contains a crystalline aluminosilicate zeolite having an SiO2/AI203 molar ratio of at least 12 and a constraint index between 1 and 12. From the aromatic hydrocarbon mixture thus obtained two fractions should then be separated, viz. an isobutanecontaining gaseous fraction, which is contacted with an alkylation catalyst and an aromatic liquid fraction boiling in the gasoline range. Finally, a fraction boiling in the gasoline range is separated from the product obtained in the alkylation, and this fraction is mixed with the gasoline fraction that was separated from the reaction product of carbon monoxide and hydrogen.

If the C2-fraction of the reaction product obtained in the conversion of the H2/CO mixture has a sufficiently high hydrocarbon content, the process may be improved by steam reforming this C2-fraction in order to prepare an additional H2/CO mixture and to mix the latter HJCO mixture with the H2/CO mixture already obtained by gasification of the coal.

The present application therefore relates to a process for preparing liquid hydrocarbons from coal, in which a) the coal is converted into a mixture of carbon monoxide and hydrogen by gasification at a temperature between 1050 and 2000"C; b) the mixture of carbon monoxide and hydrogen is converted into an aromatic hydrocarbon mixture using a catalyst which contains a crystalline aluminosilicate zeolite having an SiO2/A12O3 molar ratio of at least 12 and a constraint index between 1 and 12; c) from the aromatic hydrocarbon mixture an isobutane-containing gaseous fraction and an aromatic liquid fraction boiling in the gasoline range are separated; d) the isobutane-containing gaseous fraction is converted by alkylation into a product from which a fraction boiling in the gasoline range is separated, and e) the two fractions boiling in the gasoline range obtained according to c) and d) are mixed.

In the first step of the process according to the invention a mixture of carbon monoxide and hydrogen is prepared by gasification of coal at a temperature between 1050 and 2000"C. As a result of the use of this high temperature the synthesis gas prepared contains very little methane, if any at all. In comparison with a process in which in the first step a lower temperature is used, for instance between 800 and 1000"C, the process according to the invention gives a higher yield of CO and H2 per tonne of coal and a higher gasoline yield per tonne of coal. Because of the use of a gasification temperature between 1050 and 2000"C the product contains only very small amounts of nongaseous by-products such as tar, phenols and condensable hydrocarbons, if any at all.

The absence of these products also leads to a higher yield of CO and H2 and therefore to a higher gasoline yield than when a lower temperature is used in the gasification step.

In addition, no provisions have to be made to remove tar, phenols and condensable hydrocarbons from the synthesis gas, which will promote the economy of the gasoline preparation.

The starting materials in the process according to the invention may, for instance, be: lignite, bituminous coal, sub-bituminous coal, anthracite and coke. With a view to achieving more rapid and complete gasification, if is preferred first to reduce the starting material to powder. The high-temperature gasification is preferably carried out in the presence of oxygen and steam. It is preferred to choose such an oxygen/steam ratio that per part by volume of oxygen from 5 to l50V0v steam is present. The oxygen used is preferably preheated before it is contacted with the coal This preheating can very conveniently be carried out by heat exchange, for instance, with the hot product gas prepared according to step a) of the process. By preheating the oxygen is preferably brought to a temperature between 200 and 500"C.

The reactor in which the gasification is carried out preferably consists of an empty steel vessel lined with a heat-resistant material. A suitable reactor is described in British patent applications 18550/75 and 35133/75 (Serial Nos 1,501,284 and 1,517,765). The high temperature at which the gasification is effected is produced by the reaction of the coal with oxygen and steam. The mixture to be reacted is preferably introduced into the reactor at high speed. A suitable linear speed is 10 to 100 m/s. The pressure at which the gasification is carried out may vary within wide limits. The absolute pressure is preferably 1 to 200 bar. In order to convert as much as possible of the coal introduced into the reactor into gas, the coal particles should remain in the reactor for some time.

It has been found that a residence time of from 0.1 to 12 seconds is sufficient for this purpose. After the coal has been converted into gas, the reaction product, which consists substantially of H2, CO, CO2 and H2O, is removed from the reactor. This gas, which has as a rule a temperature higher than 1100 C, may contain impurities such as ash, carbon-containing solids and hydrogen sulphide. To allow the impurities to be removed from the gas, the latter should first be cooled. This cooling can very suitably be effected in a boiler, in which steam is formed with the aid of the waste heat.

Although as a rule the solids content of th crude gas that leaves the boiler is low, a further reduction of the solids content may nevertheless be desirable, for instance, if the gas is to be desulphurized. To this end the gas is preferably conducted through a scrubber where it is washed with water. An apparatus for this purpose is described in British patent specification 826,209. Such a washing produces a gas containing hardly any solids any more and having a temperature between 20 and 80"C. The gas may be purified still further by removal of H2S and, if desired, part of the CO2, The removal of H2S and CO2 is preferably carried out with the aid of the ADIP process or the SULFINOL process, which processes are described in British patent specifications 1,444,963, 1,131,989, 965,358, 957,260 and 972,140.

The mixture of carbon monoxide and hydrogen prepared according to the first step of the process according to the invention, is converted in the second step into an aromatic hydrocarbon mixture using a catalyst which contains a crystalline aluminosilicate zeolite of a special class.

These zeolites effect a high conversion of aliphatic hydrocarbons into aromatic hydrocarbons in commercially desirable yields and they are in general very active in conversion reactions in which aromatic hydrocarbons are involved. Although they have an uncommonly low alumina content, i.e. a high SiO2Al2O3 molar ratio, they are very active, even when the SiO2/AI203 molar ratio is more than 30. This activity is surprising because the catalytic activity of zeolites is generally ascribed to the aluminium atoms of the lattice and the cations present in combination with these aluminium atoms. These zeolites retain their crystalline character for a very long time in spite of the presence of steam, even at high temperatures such as those which effect irreversible collapse of the crystal lattice of other zeolites, e.g. those of the X- and Atype. If carbon-containing deposits are formed, they can be removed by burning them at temperatures that are higher than the temperatures usually employed for restoring the activity. In many media zeolites of this group show a very slight capability of forming coke, as a result of which the operational times between regenerations are very long.

An important property of the crystal structure of this class of zeolites is that it provides constrained access to and egress from the intracrystalline free space, because the pore size is more than about 5 A and the pore windows are of about the same size as are provided by rings of 10 oxygen atoms.

Obviously, these rings are those formed by the regular arrangement of the tetrahedrons forming the anionogenic lattice of the crystalline aluminosilicate, the oxygen atoms themselves being bound to the silicon or aluminium atoms in the centres of the tetrahedrons. In short, the zeolites that are preferably used according to the invention have a ratio of silica to alumina of at least 12 and a structure that gives constrained access to the free space in the crystals.

The said ratio of silica to alumina can be determined by usual analysis. This ratio serves the purpose of representing as precisely as possible the ratio in the rigid anionogenic lattice of the zeolite crystal, so that aluminium in the binder material or in cationogenic or other form in the channels is excluded. Although zeolites having a molar SiOjAl2O3 ratio of at least 12 are suitable, use is preferably made of zeolites having a higher ratio of at least 30 and in particular having an SiOAl2O3 ratio between 60 and 400. After activation, these zeolites obtain an intracrystalline sorptive power for n-hexane that is greater than for water, i.e. they show hydrophobic properties. It is assumed that this hydrophobic nature is an advantage in the present invention.

The zeolites that are suitable according to the invention freely sorb n-hexane and have a pore size of more than 5 A. The structure must further provide constrained access to certain large molecules. Sometimes it is possible to infer from a known crystal structure whether such a constrained access exists. If, for instance, the only pore windows in a crystal are formed by rings of eight oxygen atoms, the access for molecules having a larger cross-section than n-hexane is excluded and then the zeolite is not of the desired type. Zeolites with windows or rings with 10 atoms are preferred, although an excessive puckering or pore blockage may deactivate these zeolites. In general, zeolites with windows of rings with 12 atoms have been found to give no sufficiently constrained access to effect the conversions desired according to the invention, although as a result of pore blockage or other causes structures are possible here which are active.

Instead of trying to judge from the crystal structure whether a zeolite has the required constrained access or not, a simple constraint index determination can be carried out by continuously passing a mixture of equal quantities by weight of nhexane and 3-methylpentane at atmospheric pressure over a small sample, about 1 g or less, of the zeolite according to the process given hereinafter. A sample of the zeolite in the form of granules or extrudate is ground to a particle size which is about equal to that of coarse sand and introduced into a glass tube. Before the examination the zeolite is treated for at least 15 minutes with a stream of air of about 538"C. The zeolite is thereupon purged with helium and the temperature is set at a value between about 285"C and about 510"C to give a total conversion between 10 /n and 60%. The mixture of hydrocarbons is passed over the zeolite at a volume velocity of 1 (i.e. 1 volume of liquid hydrocarbon per volume of zeolite per hour), the mixture being diluted with helium such that the molar ratio of helium to total hydrocarbons is 4:1.

After a running rime of 20 minutes a sample of the effluent is taken and analysed (the best way is by gas chromatography) in order to determine the fraction of each of the two hydrocarbons that has not been converted.

The constraint index is calculated as follows: 10,,, (remaining fraction of n-hexane) Constraint index= 10flog (remaining fraction of 3-methylpentane) The constraint index approaches the ratio of the velocity constants for cracking the two hydrocarbons. Catalysts which are suitable for the present process are those containing a zeolite with a constraint index between 1 and 12. For some representative materials, some of which fall outside the scope of the invention the values for the constraint index (CI) are given below ZSM-5 8.3 ZSM-l 1 8.7 ZSM-12 2 ZSM-38 2 ZSM-35 4.5 TMA-offretite 3.7 Beta 0.6 ZSM-4 0.5 Acid mordenite 0.5 REY 0.4 Amorphous silica-alumina 0.6 Erionite 38 Examples of zeolites of the class defined here are ZSM-5, ZSM-l 1, ZSM-12, ZSM-35 and ZSM-38. U.S. patent 3,702,886 describes ZSM-5. ZSM-l 1 is described in U.S. patent 3,709,979 and ZSM-12 in U.S.

patent 3,832,449. U.S. patents 4,016,245 and -4,046,859 describe ZSM-35 and ZSM-38, respectively.

Naturally occurring zeolites may sometimes be converted into this type of zeolite by various activation methods and other treatments such as base exchange, steam treatment, alumina extraction and calcination or combinations of these treatments. Of the naturally occurring minerals that may be treated in this way are to be mentioned: ferrierite, brewsterite, stilbite, dachiardite, epistilbite, heulandite and clinoptilolite. The crystalline aluminosilicate zeolites which are preferably used are ZSM-5, ZSM-l 1, ZSM12, ZSM-35 and ZSM-38, particular preferance being given to ZSM-5.

According to a preferred aspect of the invention the zeolites used in the catalysts have in the dry hydrogen form a crystal lattice density of at least 1.6 g per cm3. The dry-state density can for known structures be calculated from the number of silicon plus aluminium atoms per 1000 A3, as described for instance on page 19 of the article on zeolite structure by W. U. Meier.

This article is to be found in "Proceedings of the Conference on Molecular Sieves", London, April 1967, published by the Society of Chemical Industry, London, 1968. If the crystal structure is unknown, the density of the crystal lattice can be determined according to classical pycnometer methods. The density may be determined, for instance, by immersing the zeolite in the dry hydrogen form in an organic solvent which is not sorbed by the crystal. It may be that the extraordinary, long lasting activity and stability of this class of zeolites is connected with the high density of the anionogenic lattice of the crystal, which is at least 1.6 g per cm3.

Obviously, this high density has to be associated with a relatively small free space in the crystal, which may be expected to lead to stabler structures. However, this free space seems to be important as the seat of the catalytic activity.

Below the densities are given of the crystal lattice of some representative zeolites, of which some fall outside the scope of the invention.

Volume of cavities, Density of lattice, Zeolite cm3/cm3 g/cm3 Ferrierite 0.28 1.76 Mordenite 0.28 1.7 ZSM-5, -11 0.29 1.79 Dachiardite 0.32 1.72 L 0.32 1.61 Clinoptilolite 0.34- 1.71 Laumontite 0.34 1.77 ZSM-4(omega) 0.38 1.65 Heulandite 0.39 1.69 P 0.41 1.57 Offretite 0.40 1.55 Levynite 0.40 1.54 Erionite 0.35 1.51 Gmelenite 0.44 1.46 Chabazite 0.47 1.45 A 0.5 1.3 Y 0.48 1.27 In step b) of the process according to the invention a mixture of carbon monoxide and hydrogen should be converted into an aromatic hydrocarbon mixture. Step b) may in itself be carried out as a one-step or as a two-step process. In the two-step process the mixture of carbon monoxide and hydrogen is contacted in the first step with a catalyst containing one or more metal components having catalytic activity for the conversion of a H2/CO mixture into hydrocarbons and/or oxygen-containing hydrocarbons. In the second step the product thus obtained is converted into an aromatic hydrocarbon mixture by contacting it under aromatization conditions with the crystalline aluminosilicate zeolite. In the one-step process the mixture of carbon monoxide and hydrogen is contacted with a bifunctional catalyst which contains, in addition to the crystalline aluminosilicate zeolite, one or more metal compounds having catalytic activity for the conversion of a HCO mixture into hydrocarbons and/or oxygen-containing hydrocarbons.

Step b) of the process according to the invention is preferably carried out as a onestep process.

According to step a) in the process according to the invention a H2/CO mixture is prepared, whose H2/CO molar ratio, depending on starting material and reaction conditions, may vary within wide limits.

Before this mixture is further converted according to step b) its H2/CO molar ratio can be changed by adding hydrogen or carbon monoxide. The hydrogen content of the mixture may also be increased by subjecting it to the known water gas shift reaction.

As the feed for step b) of the process according to the invention use is preferably made of a gas mixture whose HJCO molar ratio is more than 0.4. If the mixture of carbon monoxide and hydrogen used in the process according to the invention as the feed for step b) has a H2/CO molar ratio of less than 1.0 step b) is preferably carried out as a one-step process by contacting the gas with a trifunctional catalyst which contains one or more metal components having catalytic activity for the conversion of a HCO mixture into hydrocarbons and/or oxygen-containing hydrocarbons, one or more metal components having catalytic activity for the water gas shift reaction and the crystalline aluminosilicate zeolite. The ratio in which the three catalytic functions are present in the catalyst may vary within wide limits and is chiefly determined by the activity of each of the catalytic functions.

When use is made of a trifunctional catalyst in step b) of the process according to the invention for converting a H2/CO mixture with a HCO molar ratio of less than 1.0, the object is that of the acyclic hydrocarbons and/or oxygen-containing hydrocarbons formed under the influence of a first catalytic function, as much as possible is converted under the influence of a second catalytic function into an aromatic hydrocarbon mixture substantially boiling in the gasoline range, and that of the water liberated in the conversion of the mixture of carbon monoxide and hydrogen into hydrocarbons and/or in the conversion of oxygen-containing hydrocarbons into an aromatic hydrocabon mixture, as much as possible reacts under the influence of a third catalytic function with the carbon monoxide present in an excess amount in the mixture of carbon monoxide and hydrogen with formation of a mixture of hydrogen and carbon dioxide. In the composition of an optimum trifunctional catalyst to be used in step b) of the process according to the invention, which catalyst contains a given quantity of a first catalytic function having a given activity, it is therefore possible to do with less of the other catalytic functions according as these are more active.

Although the trifunctional catalysts that can be used in step b) of the process according to the invention are described in this patent application as catalysts containing one or more metal components having catalytic activity for the conversion of a HJCO mixture into hydrocarbons and one or more metal components having catalytic activity for the water gas shift reaction, this means in no way that metal components each having in themselves one of the two catalytic functions should always separately be present in the catalysts. For, it has been found that metal components and combinations of metal components having catalytic activity for the conversion of a HJCO mixture into substantially oxygencontaining hydrocarbons as a rule also have sufficient catalytic activity for the water gas shift reaction so that in such a case incorporation of one metal component or one combination of metal components into the catalysts will suffice. Examples of such metal components are the metals chosen from the group formed by the metals zinc, copper and chromium. When use is made of trifunctional catalysts containing these metals in step b) of the process according to the invention, preference is given to catalysts containing combinations of at least two of these metals, for instance, the combination zinc-copper, zinc chromium, or zinc-copper-chromium.

Particular preference is given to a trifunctional catalyst containing in addition to the crystalline aluminosilicate zeolite the metal combination zinc-chromium. Metal components and combinations of metal components having catalytic activity for the conversion of a HJCO mixture into substantially hydrocarbons have as a rule no or insufficient activity for the water gas shift reaction. When use is made of such metal components or combinations of metal components in the catalyst, one or more separate metal components having catalytic activity for the water gas shift reaction should therefore be incorporated therein.

The trifunctional catalysts which are used in step b) of the process according to the invention are preferably composed of two or three separate catalysts, which will for convenience be designated catalysts X, Y and Z. Catalyst X is the catalyst containing the metal components having catalytic activity for the conversion of a HJCO mixture into hydrocarbons and/or oxygencontaining hydrocarbons. Catalyst Y is the crystalline alumino-silicate zeolite. Catalyst Z is the catalyst containing the metal component having catalytic activity for the water gas shift reaction. As has been explained hereinbefore the use of Z-catalyst may be omitted in some cases.

If as the X catalyst a catalyst is used which is capable of converting a HJCO mixture into substantially oxygencontaining hydrocarbons, preference is given to a catalyst which is capable of converting the HJCO mixture into substantially methanol and/or dimethyl ether. For the conversion of a HJCO mixture into substantially methanol, catalysts containing the metal combinations mentioned hereinbefore are very suitable. If desired the said metal combinations may be emplaced on a carrier material. By introducing an acid function into these catalysts, for instance by emplacing the metal combination on an acid carrier, it may be effected that apart from the conversion of the HJCO mixture into methanol a considerable part of the mixture will be converted into dimethyl ether.

X-catalysts which are capable of converting a H2/CO mixture into substantially hydrocarbons are referred to in the literature as Fischer-Tropsch catalysts. Such catalysts often contain one or more metals of the iron group or ruthenium together with one or more promotors to increase the activity and/or selectivity and sometimes a carrier material such as kieselguhr. They can be prepared by precipitation, melting and by impregnation.

The preparation of the catalysts containing one or more metals of the iron group, by impregnation, takes place by impregnating a porous carrier with one or more aqueous solutions of salts of metals of the iron group and, optionally, of promotors, followed by drying and calcining the composition. If in step b) of the process according to the invention use is made of a catalyst combination in which catalyst X is a Fischer-Tropsch catalyst, it is preferred to choose for this purpose an iron or cobalt catalyst, in particular such a catalyst which has been prepared by impregnation. Very suitable Fischer-Tropsch catalysts for use in the catalyst combinations according to the invention are the catalysts prepared by impregnation according to the British patent application No. 46427/77 (Serial No.

1,548,468). The catalysts concerned contain per 100 pbw carrier 10--75 pbw of one or more metals of the iron group, together with one or more promotors in a quantity of 1 50 /n of the quantity of metals of the iron group present on the catalyst, which catalysts have such a specific average pore diameter (p) of at most 10,000 nm and such a specific average particle diameter (d) of at most 5 mm, that the quotient p/d is more than 2 (p in nm and d in mm).

If in step b) of the process according to the invention the object is to use a catalyst combination of which X is a Fischer Tropsch iron catalyst, it is preferred to choose an iron catalyst containing a promotor combination consisting of an alkali metal, a metal that is easy to reduce, such as copper or silver and, optionally, a metal that is hard to reduce, such as aluminium or zinc. A very suitable iron catalyst for the present purpose is a catalyst prepared by impregnation containing iron, potassium and copper on silica as the carrier. If in step b) of the process according to the invention the object is to use a catalyst combination of which X is a Fischer-Tropsch cobalt catalyst, it is preferred to choose a cobalt catalyst containing a promotor combination consisting of an alkaline-earth metal and thorium, uranium or cerium. A very suitable Fischer-Tropsch cobalt catalyst for the present purpose is a catalyst prepared by impregnation containing cobalt, magnesium and thorium on silica as the carrier. Other very suitable Fischer-Tropsch cobalt catalysts prepared by impregnation are catalysts containing, in addition to cobalt, one of the elements chromium, titanium, zirconium and zinc on silica as the carrier. If desired, it is also possible to use in step b) of the process according to the invention catalyst combinations containing an X-catalyst, which is capable of converting a H2/CO mixture into a mixture containing both hydrocarbons and oxygen-containing hhydrocarbons in comparable quantities.

As a rule, such a catalyst has sufficient catalytic activity for the water gas shift reaction, so that the use of a Z-catalyst in the combination can be omitted. An example of an X-catalyst of this type is an iron-chromium oxide catalyst. If desired. it is also possible to use in step b) of the process according to the invention catalyst combinations containing two or more Xcatalysts, for instance in addition to a catalyst of the X-type which is capable of converting a H;CO mixture into substantially hydrocarbons, a second catalyst of the X-type which is capable of converting a H2/CO mixture into substantially oxygen-containing hydrocarbons.

Z-catalysts which are capable of converting a H2/CO mixture into a HJCO2 mixture are referred to in the literature as CO shift catalysts. Such catalysts often contain one or more metals of the group formed by iron, chromium, copper, zinc, cobalt, nickel and molybdenum as the catalytically active component, either as such, or in the form of their oxides or sulphides. Examples of suitable CO-shift catalysts are the may also be carried out in fluidized-bed operation or with the use of a suspension of the catalyst hydrocarbon oil. It is preferred to use catalyst particles then with a diameter between 10 and 150 mm.

In the process according to the invention an isobutane-containing gaseous fraction and an aromatic liquid fraction boiling in the gasoline range should be separated from aromatic hydrocarbon mixture obtained according to step b). It is preferred to separate the reaction mixture originating from step b) in step c) into a C2~fraction, a propane fraction, an isobutane-containing fraction, an n-butane fraction and an aromatic liquid fraction boiling in the gasoline range. The C2-fraction may be used as fuel gas. If desired, a HiCO mixture can be separated from the C2-fraction, which mixture may be recirculated to step b). If the hydrocarbon content of the C;fraction is sufficiently high, it may be preferred to subject it, either after removal of a HCO mixture from it or not, to steam reforming in order to prepare additional synthesis gas, which may be used as feed component for step b). Steam reforming of the C2-fraction can very suitably be effected by contacting it together with steam at elevated temperature and pressure with a nickelcontaining catalyst. Water which may be formed as a by-product in step b) may, if desired, be used in the process in the steam gasification of the coal and/or in the steam reforming of the C;fraction.

In step d) of the process according to the invention the isobutane-containing gaseous fraction should be converted by alkylation into a product from which a fraction boiling in the gasoline range can be separated. This alkylation can very conveniently be effected by contacting the fraction with a strong acid as the catalyst, such as sulphuric acid or hydrofluoric acid. Since the gaseous part of the reaction product of step b) usually contains only small amounts olefins, the isobutane-containing gaseous fraction which is separated from it will often have too low an olefin content to realize a sufficient conversion of the isobutane present in it by alkylation. It is therefore preferred to increase the olefin content of the fraction before subjecting it to alkylation. An increase in the olefin content of the isobutane-containing fraction can conveniently be effected by mixing it with an olefin-rich stream which may originate from an external source or which has been prepared by dehydrogenation of the paraffins obtained in the process, such as a propane fraction, an n-butane-fraction or an LPG fraction obtained from it by mixing.

Dehydrogenation of these fractions can conveniently be effected by contacting them at elevated temperature with a chromium-containing catalyst. From the product obtained in the alkylation a fraction boiling in the gasoline range is separated and this fraction is mixed according to step e) of the process according to the invention with the aromatic liquid fraction obtained in step c) and boiling in the gasoline range the non-converted isobutane is preferably separated from the product obtained in the alkylation and recirculated to the alkylation reaction. In order to increase the vapour pressure of the gasoline mixture thus obtained, light hydrocarbons are preferably added to it. As light hydrocarbons use can very conveniently be made of n-butane or LPG, which may be obtained as by-products of the process.

Two process schemes for the conversion of coal into aromatic gasoline according to the invention will be explained in more detail hereinafter with the aid of the figures.

Process scheme I (see Fig. I) The process is carried out in an apparatus comprising successively a gasification section (1), a gas purification section (2), a hydrocarbon synthesis section (3), the first separation section (4) a steam reforming section (5), an alkylation section (6) and the second separation section (7). A mixture of coal (8), oxygen (9) and steam (10) is gasified and the crude gas (11) is purified. The purified gas (12) is mixed with a synthesis gas (13) reverted to further on and prepared by steam reforming, and the mixture is converted under the influence of a trifunctional catalyst according to the invention into an aromatic hydrocarbon mixture (14). This hydrocarbon mixture is separated into a C7fraction (15), a propane fraction (16) an isobutane fraction (17), an n-butane fraction (18) and an aromatic gasoline fraction (19). the C2-fraction (15) is converted by steam reforming into the synthesis gas (13). The propane fraction (16) and the n-butane fraction (18) are mixed into the LPG fraction (20), which is subsequently separated into two portions (21) and (22) having the same composition.

The isobutane fraction (17) is alkylated together with an olefin stream (23) originating from an external source and with an isobutane recirculation stream (24) reverted to later. From the alkylation product (25) the isobutane recirculation stream (24) and a gasoline fraction (26) are separated. The gasoline fraction (26) is mixed with the gasoline fraction (19) and with portion (22) of the LPG fraction (20) into gasoline (27).

Process Scheme II (see Fig. 2) The process is carried out in an apparatus comprising successively a gasification section (1), a gas purification section (2), a hydrocarbon synthesis section (3), the first separation section (4), a dehydrogenation section (5), an alkylation section (6) and the second separation section (7). A mixture of coal (8), oxygen (9) and steam (10) is gasified, the crude gas (11) is purified and the purified gas (12) is converted under the influence of a trifunctional catalyst according to the invention into an aromatic hydrocarbon mixture (13). This hydrocarbon mixture is separated into a C2-fraction (14), a propane fraction (15), an isobutane fraction (16) an n-butane fraction (17) and an aromatic gasoline fraction (18).

The propane fraction (15) is separated into two portions (19) and (20). Portion (20) is converted by dehydrogenation into a mixture of propene and propane (21). The n-butane fraction (17) is separated into two portions (22) and (23). The isobutane fraction (16) is alkylated together with the propane/propene stream (21) and with an isobutane recirculation stream (24) reverted to later. From the alkylated product (25) a propane fraction (26), the isobutane recirculation stream (24) and a gasoline fraction (27) are separated. The propane fraction (26) is mixed with portion (19) of the propane fraction (15) and with portion (23) of the n-butane fraction (17) into the LPG fraction (28). The gasoline fraction (27) is mixed with the gasoline fraction (18) and with portion (22) of n-butane fraction (17) into the gasoline (29).

The present invention also comprises equipment for carrying out the process accordng to the invention as shown schematically in Figures 1 and 2.

The invention will now be further explained with the aid of the following examples.

Example I ZSM-5 (zeolite A) was prepared as follows. A mixture of SiO2, NaAIO2, NaOH and [(C3H,)4N]OH in water with the molar composition 1.21 Na2O . A12O3 .9[(C,H,),N],O . 9[(C3H7)4N]2O.

29.1 SiO2 .480 H2O was heated for 98 hours in an autoclave at 1500C under autogeneous pressure. After having cooled the reaction mixture, the zeolite formed was filtered off, washed with water until the pH of the wash water was about 8 and dried for two hours at 1200C.

With zeolite A as the starting material zeolite B was prepared by, successively, calcining zeolite A at 5000C, boiling with 1.0 molar NH4NO3 solution, washing with water, boiling again with 1.0 molar NH4NO3 solution and washing, drying for two hours at 1200C and calcining for four hours at 500"C.

Example II A catalyst C was prepared by thoroughly mixing equal parts by weight of the following finely powdered materials a) a Fe/Cu/K/SiO2 Fishcher-Tropsch catalyst prepared by impregnation; b) zeolite B; c) a Cu/Zn CO shift catalyst.

Catalyst C was extruded to particles having a diameter of 1-3 mm.

Example III A catalyst D was prepared by mixing a ZnO-Cr2O composition with zeolite B in a weight ratio of 5:1. Both materials were present in the catalyst in the form of particles having a diameter of 0.15-0.3 mm. The ZnO-Cr2O3 composition used catalyses both the reduction of CO to methanol and the water gas shift reaction.

Example IV Bituminous coal was gound to a particle size of less than 120 microns and used as the feed for a high-temperature coal gasifier.

Per gramme of coal 0.9 g oxygen and 0.15 g steam were added. The coal gasification was effected at a temperature of 1500"C, a pressure of 30 bar and a residence time of 0.5 s. The coal conversion was 99%. The gas obtained had the following composition v CH4 0.1 CO 64.7 H2 31.8 CO2 1.7 The gas further contained about 1.7 ,Xv H2O, COS and H2S. To remove the lastmentioned impurities from the gas, this gas was passed at about 45"C through a mixture of diisopropyl amine, sulfolane and water.

The resulting synthesis gas, of which the CO/H2 molar ratio was 2.03, was further purified by passing it at 2000C over ZnO.

The synthesis gas thus purified was used in Examples V and VI which were carried out according to process schemes I and II, respectively.

Example V The synthesis gas prepared according to Example IV was mixed in a volume ratio of 60:40 with a synthesis gas (If), reverted to further on, obtained by steam reforming and the mixture was contacted with catalyst C at a temperature of 2800 C, a pressure of 30 bar and a space velocity of 1000 1.1-' . h-'. The synthesis gas conversion was 85"/,. The hydrogen mixture obtained had the following composition %w C, 14 C2 13 C3 14 n-C4 6 iC4 8 Csfgasoline 45 The olefin content of both the C3 and the C4 fractions was less than 1%w. The reaction product was separated by cooling into a C2-fraction (including carbon dioxide and unconverted synthesis gas) and a C3+fraction. The C2-fraction was mixed with 25 Ev steam and the mixture was converted into synthesis gas by contacting it at a temperature of 900"C and a pressure of 30 bar with a Ni-containing catalyst. The product was washed with caustic solution to remove CO2 and the purified synthesis gas (II) was mixed with the feed gas. The C3+fraction was separated into a propane fraction, an isobutane fraction, an n-butane fraction and a Cs+gasoline fraction. The propane fraction and the n-butane fraction were mixed into LPG. The isobutane fraction was mixed with 80%v of a C3-C5 olefin mixture originating from an external source and the mixture was converted by contacting it at 400C with a HF alkylation catalyst. By recirculation of isobutane a constant isobutane/olefin ratio of 14 was maintained in the alkylation reactor. The alkylate which was obtained in 95% yield was mixed with the gasoline obtained earlier in the process. To bring the vapour pressure of the mixture to the proper value part of the LPG was added to it. The gasoline thus obtained had an octane number (CRON) of 91.

Example VI to The s.1thesis gas prepared according Example IV was contacted at a temperature of 375"C, a pressure of 60 bar and a space velocity of 300 I.l-'.h-' with catalyst D.

The synthesis gas conversion was 95%. The reaction developed completely to carbon dioxide. The hydrocarbon mixture obtained had the following composition: w Cg 4 C2 6 C3 39 n-C4 3 iC4 8 Cs+gasoline 42 The olefin content of both the C3 and the C4 fractions was less than 10/,w. The reaction product was separated by cooling into a C2-fraction (including carbon dioxide and unconverted synthesis gas) and a C3+fraction was separated into a propane fraction, an isobutane fraction, an n-butane fraction and a Cs+gasoline fraction mainly consisting of aromatics. The propane fraction was divided into two equal portions of which one was converted by dehydrogenation at 6000C over a Cr2O3 catalyst into a mixture of propane and propene. The convertion from propane into propene was 30%. The propane/propene mixture thus obtained was mixed with the isobutane fraction and the mixture was converted by contacting it at 400C with a HF alkylation catalyst. From the product obtained in the alkylation a propane fraction, an isobutane fraction and a gasoline fraction was separated. By recirculation of isobutane a constant isobutanelolefin ratio of 14 was maintained.

The alkylation gasoline yield was 94%. The alkylation gasoline was mixed with the gasoline obtained earlier in the process. To bring the vapour pressure of the mixture to the proper value, part of the n-butane fraction was added. The gasoline thus obtained had an octane number (CRON) of 96. The remaining part of both the propane fraction and the n-butane fraction obtained from the C3+fraction of the hydrocarbon synthesis product were mixed with the propane fraction obtained from the alkylation product, into LPG.

WHAT WE CLAIM IS:- 1. A process for preparing liquid hydrocarbons from coal, characterized in that a) the coal is converted into a mixture of carbon monoxide and hydrogen by gasification at a temperature between 150 and 2000"C; b) the mixture of carbon monoxide and hydrogen is converted into an aromatic hydrocarbon mixture using a catalyst which contains a crystalline aluminosilicate zeolite having an SiO2/Al2O3 molar ratio of at least 12 and an constraint index between 1 and 12; c) from the aromatic hydrocarbon mixture an isobutane-containing gaseous fraction and an aromatic liquid fraction boiling in the gasoline range are separated; d) the isobutane-containing gaseous fraction is converted by alkylation into a product from which a fraction boiling in the gasoline range is separated, and e) the two fractions boiling in the gasoline range obtained according to c) and d) are mixed.

2. A process as claimed in claim 1, characterized in that the gasification is

**WARNING** end of DESC field may overlap start of CLMS **.

Claims (11)

**WARNING** start of CLMS field may overlap end of DESC **. %w C, 14 C2 13 C3 14 n-C4 6 iC4 8 Csfgasoline 45 The olefin content of both the C3 and the C4 fractions was less than 1%w. The reaction product was separated by cooling into a C2-fraction (including carbon dioxide and unconverted synthesis gas) and a C3+fraction. The C2-fraction was mixed with 25 Ev steam and the mixture was converted into synthesis gas by contacting it at a temperature of 900"C and a pressure of 30 bar with a Ni-containing catalyst. The product was washed with caustic solution to remove CO2 and the purified synthesis gas (II) was mixed with the feed gas. The C3+fraction was separated into a propane fraction, an isobutane fraction, an n-butane fraction and a Cs+gasoline fraction. The propane fraction and the n-butane fraction were mixed into LPG. The isobutane fraction was mixed with 80%v of a C3-C5 olefin mixture originating from an external source and the mixture was converted by contacting it at 400C with a HF alkylation catalyst. By recirculation of isobutane a constant isobutane/olefin ratio of 14 was maintained in the alkylation reactor. The alkylate which was obtained in 95% yield was mixed with the gasoline obtained earlier in the process. To bring the vapour pressure of the mixture to the proper value part of the LPG was added to it. The gasoline thus obtained had an octane number (CRON) of 91. Example VI to The s.1thesis gas prepared according Example IV was contacted at a temperature of 375"C, a pressure of 60 bar and a space velocity of 300 I.l-'.h-' with catalyst D. The synthesis gas conversion was 95%. The reaction developed completely to carbon dioxide. The hydrocarbon mixture obtained had the following composition: w Cg 4 C2 6 C3 39 n-C4 3 iC4 8 Cs+gasoline 42 The olefin content of both the C3 and the C4 fractions was less than 10/,w. The reaction product was separated by cooling into a C2-fraction (including carbon dioxide and unconverted synthesis gas) and a C3+fraction was separated into a propane fraction, an isobutane fraction, an n-butane fraction and a Cs+gasoline fraction mainly consisting of aromatics. The propane fraction was divided into two equal portions of which one was converted by dehydrogenation at 6000C over a Cr2O3 catalyst into a mixture of propane and propene. The convertion from propane into propene was 30%. The propane/propene mixture thus obtained was mixed with the isobutane fraction and the mixture was converted by contacting it at 400C with a HF alkylation catalyst. From the product obtained in the alkylation a propane fraction, an isobutane fraction and a gasoline fraction was separated. By recirculation of isobutane a constant isobutanelolefin ratio of 14 was maintained. The alkylation gasoline yield was 94%. The alkylation gasoline was mixed with the gasoline obtained earlier in the process. To bring the vapour pressure of the mixture to the proper value, part of the n-butane fraction was added. The gasoline thus obtained had an octane number (CRON) of 96. The remaining part of both the propane fraction and the n-butane fraction obtained from the C3+fraction of the hydrocarbon synthesis product were mixed with the propane fraction obtained from the alkylation product, into LPG. WHAT WE CLAIM IS:-

1. A process for preparing liquid hydrocarbons from coal, characterized in that a) the coal is converted into a mixture of carbon monoxide and hydrogen by gasification at a temperature between 150 and 2000"C; b) the mixture of carbon monoxide and hydrogen is converted into an aromatic hydrocarbon mixture using a catalyst which contains a crystalline aluminosilicate zeolite having an SiO2/Al2O3 molar ratio of at least 12 and an constraint index between 1 and 12; c) from the aromatic hydrocarbon mixture an isobutane-containing gaseous fraction and an aromatic liquid fraction boiling in the gasoline range are separated; d) the isobutane-containing gaseous fraction is converted by alkylation into a product from which a fraction boiling in the gasoline range is separated, and e) the two fractions boiling in the gasoline range obtained according to c) and d) are mixed.

2. A process as claimed in claim 1, characterized in that the gasification is

carried out in the presence of oxygen and steam.

3. A process as claimed in claim 2, characterized in that a coaVoxygen/steam mixture is introduced into a gasification reactor at a linear speed from 10 to 100 m/s.

4. A process as claimed in any one of the preceding claims, characterized in that the crystalline aluminosilicate has a SiO2/AI203 molar ratio between 60 and 400.

5. A process as claimed in any one of the preceding claims, characterized in that a trifunctional catalyst is used in step b).

6. A process as claimed in any one of the preceding claims characterized in that step b) is carried out at a temperature from 200 to 5000C, a pressure from 1 to 150 bar and a space velocity from 50 to 5000 Nl gas/l catalyst/hour.

7. A process as claimed in any one of the preceding claims characterized in that the reaction mixture originating from step b) is separated in step c) into a C2-fraction, a propane fraction, an isobutane-containing fraction, a n-butane fraction and an aromatic liquid fraction boiling in the gasoline range.

8. A process as claimed in any one of the preceding claims, characterized in that the olefin content of the isobutane-containing fraction obtained in step c) is increased before subjecting it to alkylation in step d).

9. A process as claimed in claim 1, substantially as hereinbefore described with special reference to Figures 1 and 2 of the accompanying drawings.

10. A process as claimed in claim 1 substantially as hereinbefore described with special reference to the Examples.

11. Liquid hydrocarbons whenever prepared with the aid of the process as claimed in any one of claims 1--10.