WO2014174107A1

WO2014174107A1 - Production of hydrocarbons from synthesis gas

Info

Publication number: WO2014174107A1
Application number: PCT/EP2014/058529
Authority: WO
Inventors: Qingjie Ge; Chun Wang; Xiangang MA; Hengyong Xu
Original assignee: Dalian Institute of Chemical Physics of CAS; BP PLC
Current assignee: Dalian Institute of Chemical Physics of CAS; BP PLC
Priority date: 2013-04-26
Filing date: 2014-04-25
Publication date: 2014-10-30
Anticipated expiration: 2015-10-26

Abstract

A process is described for synthesis of saturated hydrocarbons from syngas. The process contains a one stage catalyst system and/or multi-stage catalyst system for use in the catalysed production of saturated hydrocarbons from carbon oxides and hydrogen. The process decreases amounts of undesirable low carbon number hydrocarbons, such as C₃ and lower, and provides high octane number hydrocarbons boiling in the range of gasoline.

Description

■PRODUCTION OF HYDROCARBONS FROM SYNTHESIS GAS

This invention relates to the production of hydrocarbons from synthesis gas.

Examples of the invention relate to the production of saturated hydrocarbons from synthesis gas. Aspects of the invention relate to the production of saturated C₅ and higher hydrocarbons, in some examples in addition to the production of aromatic hydrocarbons for example being C9 or higher hydrocarbons. Aspects of the invention relate to the co- production of saturated C₅ to C₈ hydrocarbons and C9 - Cn aromatic hydrocarbons. In examples of the invention, the saturated C₅ and higher hydrocarbons include branched alkanes (iso-paraffins). In examples of the invention, a two-stage catalyst bed reaction scheme is used.

In recent years, the dominance of natural gas and petroleum as feedstocks has begun to be challenged. New feedstocks such as tar sands, coal, biomass and municipal waste have been increasing in importance. A move to increase the diversity of feedstocks has driven the development of synthesis gas (syngas) routes to supplement or replace conventional routes to hydrocarbons from natural gas and petroleum.

The Fischer-Tropsch (FT) process for conversion of syngas to hydrocarbons was first developed in the 1920s. The catalysts used in such processes conventionally contain active metals of the transition elements group VIII. However, in order to yield hydrocarbons with five or more carbon atoms, it has been identified that the presence of metals such as Fe, Co, Ni and Ru is desirable. The distribution of the produced hydrocarbons when these traditional FT catalysts are used follows the Anderson-Schulz-Flory law (ASF), according to which some types of product within a narrow range of number of carbon atoms cannot be obtained with high selectivity.

More recently, a so-called "modified FT process" was proposed, in which a traditional FT catalyst is supported or mixed with zeolites. Such a modified catalyst has been seen in some cases to exhibit a higher selectivity for the C₅-C₁₂ hydrocarbon product fraction. For such modified FT process, it has been suggested that the shape-selectivity of the zeolites plays a key role in hydrocarbon product distribution; Jincan Kang, et al.

Angewandte Chemie Internation Edition, 2011, 50, 5200-5203, describes the catalytic performance of the meso-ZSM-5 -supported ruthenium catalysts. Jun Bao, et al. Applied Catalysis A: General, 2011, 394, 195-200, reports the performance of a type of core-shell Fe-ZSM-5 catalyst. Both report high gasoline selectivity and high ratio of branched to linear molecules in the gasoline fraction. However, it was reported that the selectivity of methane (CH₄) was relatively high, which is generally considered to be disadvantageous.

Processes exist for selectively converting syngas to methanol. The conversion of methanol to hydrocarbons is also known, including the methanol to olefins (MTO) process, methanol to gasoline (MTG), and methanol to aromatics (MTA) processes. In such processes, the methanol is first synthesized and separated from the other components before being fed to the methanol conversion process.

There have been some proposals for the co-production of light iso-paraffins (C₅-C₈) and heavy aromatics (C9-C11) from syngas via methanol and/or dimethyl ether (DME).

In the co-production of light iso-paraffins (C₅-C₈) and heavy aromatics (C₉-Cn) from syngas, this may be achieved in a single catalyst bed reaction system. The catalyst bed may be composed of methanol synthesis catalyst and metal-modified zeolite. Alternatively, co- production of light iso-paraffins (C₅-C ) and heavy aromatics (C₉-Cn) from syngas can be realized in a two catalyst bed reaction system. The catalyst may be preferentially active to produce methanol in the first catalyst bed. Thus the catalyst in the first catalyst bed may include a methanol synthesis catalyst. The intermediate product may therefore include methanol. The catalyst may produce dimethyl ether (DME) in the first catalyst bed. In some examples, both methanol and DME are produced in the first stage. Thus the intermediate product stream may include DME and/or methanol. The catalyst of the second catalyst bed preferentially includes a dehydration catalyst. In some examples, the catalyst might be modified by some metals in the second catalyst bed. The second catalyst bed may have dehydration and hydrogenation activity. Aspects and features of the present invention can be applied as appropriate to a single catalyst bed system or to a two, or more, bed system.

Aspects of the present invention seek to provide a method for the production of hydrocarbons from syngas. In examples, the invention seeks to provide high selectivity to branched alkanes (iso-paraffins), in particular C₅₊ alkanes. In examples, the invention seeks to provide selectivity to heavy aromatics, in particular C₉₊ aromatics. In examples, the invention seeks to provide low selectivity to C₅ to C₈ aromatics.

According to an aspect of the invention there is provided a catalyst composition for use as a dehydration/ hydrogenation catalyst in a multi-stage catalyst system for the catalysed production of saturated hydrocarbons from carbon oxide(s) and hydrogen, the catalyst composition comprising: an acidic substrate comprising a zeolite, or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite or the zeolite of the M-zeolite catalyst is 100 or more.

It has now been identified that by use of a zeolite composition having a high

Si0₂/Al₂0₃ ratio can be beneficial, for example in the production of a product stream including C₅ - C₈ having a high iso-paraffinic content as well as a low olefin and aromatic content. The heavy C₉ - Cn fraction contains mainly trimethylbenzenes and

tetramethylbenzenes in some examples.

Preferably the Si0₂/Al₂0₃ molar ratio of the zeolite or M-zeolite catalyst is 100 or more, or 120 or more, or 140 or more, or 200 or more, or 250 or more, or 300 or more.

Without wishing to be bound by any particular theory, it is thought that the zeolite composition having a high Si0₂/Al₂0₃ ratio has a reduced number of acid sites compared with conventional, lower Si0₂/Al₂0₃ ratio, zeolites. It is thought that this reduced number of acid sites in the second stage catalyst improves the selectivity to iso-paraffins in the formation of the hydrocarbon product.

The Si0₂ Al₂0₃ molar ratio of the zeolite of the catalyst composition, for example a second stage catalyst composition may be 120 or more, or 140 or more. In some examples, the Si0₂/Al₂0₃ molar ratio may be 200 or more, or 250 or more, or 300 or more, or 350 or more. In examples given herein, the Si0₂/Al₂0₃ molar ratio is 360.

The Si0₂/Al₂0₃ molar ratio of the zeolites of the second stage catalyst composition of the present invention is significantly higher than for conventional zeolites which may have for example a Si0₂/Al₂0₃ molar ratio of about 20.

The Si0₂/Al₂0₃ ratio of a particular zeolite sample may be measured by any appropriate method, for example by a ICP-MS technique (inductively coupled plasma mass spectrometry),, or XRF technique (X-ray Fluroscene technique).

Where reference is made to dehydration/ hydrogenation, for example to a component or catalyst being used for dehydration/ hydrogenation, preferably it will be understood that the reference is to dehydration, hydrogenation or to both dehydration and hydrogenation as appropriate in the context.

The catalyst composition may have been prepared for example by a method described herein. However, some aspects of the invention extend to the case in which is obtained by other methods or from other sources. Thus aspects of the invention extend to such catalyst compositions irrespective of their source or method of preparation.

The zeolite catalyst may comprise one or more from the group comprising Y zeolite, β zeolite, and ZSM-5. The acidic substrate may comprise two or more such components from the group. The zeolite catalyst may comprise ZSM-5.

Preferably M comprises a hydrogenation metal. M preferably comprises a metal chosen from the group comprising Pd, Pt, Rh, Ru, Cu and Zn, preferably M comprises Pd and/or Cu.

The weight percent of metal M in the M-zeolite may be for example from about 0.1 wt% to about 20 wt%. In examples, the weight percent M in the M-zeolite is from about 0.1 wt% to about 2wt%, for example from about 0.5 wt% to about 1 wt%.

In some applications of aspects of the invention, the hydrogenation catalyst is used in combination with an additional catalyst, for example a carbon oxide(s) conversion catalyst. Thus aspects of the invention provide a catalyst system including the catalyst and a carbon oxide(s) conversion catalyst. The catalyst system may comprise a two-stage catalyst system, for example in which the two stages of the system are separate. The two-stage catalyst system may be a part of a multi-stage catalyst system. In other examples, the two catalysts may be mixed together, or may be provided having a direct interface between them, or spaced apart by a spacer element. The two catalyst stages may be separate.

Thus a further aspect of the invention provides a multi-stage catalyst system for use in the catalysed production of saturated hydrocarbons from carbon oxides and hydrogen, the catalyst system comprising a first stage comprising a carbon oxide(s) conversion catalyst, and a second stage comprising a hydrogenation catalyst comprising: an acidic substrate comprising a zeolite, or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite or of the zeolite of the M-zeolite catalyst is 100 or more.

The multi-stage catalyst system may be used as physically separate stages, or physically segmented stages, or the stages may be physically mixed, and other options are possible.

The carbon oxides conversion catalyst may be active to produce methanol and/or may be active to produce dimethyl ether (DME), for example to produce DME in the first stage where a two-stage or multi-stage system is used, or for a hybrid catalyst, to produce DME in the catalysed conversion process. In some examples, both methanol and DME may be produced in the process.

The production of methanol from carbon oxide(s) and hydrogen is equilibrium limited. The production of DME direct from carbon oxide(s) and hydrogen is less equilibrium limited. Pressure can be used to increase the yield, as the reaction which produces methanol exhibits a decrease in volume, as disclosed in US Patent No. 3,326,956.

Improved catalysts have allowed viable rates of methanol formation to be achieved at relatively low reaction temperatures, and hence allow commercial operation at lower reaction pressures. For example a CuO/ZnO/Al₂0₃ conversion catalyst may be operated at a nominal pressure of 5-10 MPa and at temperatures ranging from approximately 150 degrees C to 300 degrees C. However, at higher reaction temperatures, reduction in catalyst lifetime has commercially been found to be a problem. A low-pressure, copper- based methanol synthesis catalyst is commercially available from suppliers such as BASF and Haldor-Topsoe. Methanol yields from copper-based catalysts are generally over 99.5% of the converted carbon oxide(s) present. Water is a by-product of the conversion of C0₂ to methanol and the conversion of synthesis gas to C₂ and C₂₊ oxygenates. In the presence of an active water gas-shift catalyst, such as a methanol catalyst or a cobalt molybdenum catalyst, the water equilibrates with the carbon monoxide to give C0₂ and hydrogen.

Recently, to seek to overcome the equilibrium limitation of the methanol synthesis catalyst, direct syngas-to-DME processes have been developed. These processes are thought to proceed via a methanol intermediate which is etherified by an added acid functionality in the catalyst, for example as described in PS Sai Prasad, et al., Fuel Processing Technology Volume 89, Issue 12, December 2008, p 1281-1286.

The carbon oxide(s) conversion catalyst may be provided together with the dehydration/hydrogenation catalyst in a mixed catalyst. Thus a methanol synthesis catalyst, and/or DME synthesis catalyst and the dehydration/hydrogenation catalyst will be present together in a mixed catalyst. The mixed catalyst may for example include a mechanical mixture of the hydrogenation catalyst and a methanol synthesis catalyst and/or DME synthesis catalyst.

Therefore, a further aspect of the invention provides a mixed catalyst for the catalysed production of saturated hydrocarbons from carbon oxides and hydrogen, mixed catalyst including: a carbon oxide(s) conversion catalyst, and a hydrogenation catalyst comprising an acidic substrate comprising a zeolite, or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂/Al₂03 molar ratio of the zeolite or the zeolite of the M-zeolite catalyst is 100 or more.

The carbon oxide(s) conversion catalyst comprises a methanol synthesis catalyst and/or DME synthesis catalyst. The carbon oxide(s) conversion catalyst may comprise a copper oxide. The carbon oxide(s) conversion catalyst may comprise a zeolite and/or γ-Α1- ₂0₃.

The carbon oxide(s) conversion catalyst may for example comprise a methanol synthesis catalyst. The methanol synthesis catalyst may be any appropriate composition. In preferred examples, the catalyst includes Cu-ZnO-[Sup], Pd-[Sup] and Zn-Cr-[Sup], where [Sup] is preferably a support composition for example including A1₂0₃, Si0₂, and/or zeolite. Preferably the carbon oxide(s) conversion catalyst comprises Cu-ZnO-Al₂0₃.

The carbon oxides(s) conversion catalyst may for example comprise a DME synthesis catalyst. The DME synthesis catalyst may comprise for example Cu-ZnO [Sup], Cu-Pd/Ce0₂-[Sup], where [Sup] may be as described above, together with for example ZSM-5 or γ-Α1₂0₃ and/or ZSM-5. Preferably [Sup] comprises A1₂0₃ for the DME synthesis catalyst. Thus the carbon oxide(s) conversion catalyst may comprise Cu-ZnO- Al₂0₃/ZSM-5 Cu-ZnO-Al₂0₃/y-Al₂0_3. Thus the carbon oxide(s) conversion catalyst may comprise a hybrid catalyst.

The hybrid catalyst may be prepared by any appropriate method.

In some examples, the first stage or carbon oxide(s) conversion catalyst composition may also comprise a zeolite. It has been identified by the inventors that the Si0₂/Al₂0₃ molar ratio of zeolite of the first stage catalyst composition is of less importance in some examples than for that of the second stage catalyst. Thus in some examples, the

Si0₂/Al₂0₃ molar ratio of zeolite of the first stage catalyst composition is less than that of the zeolite of the second stage composition. For example, the Si0₂/Al₂0₃ molar ratio of zeolite of the first stage composition may be less than 100, for example less than 70, for example less than 50.

Also provided by the invention is a catalyst for use in the catalysed production of saturated hydrocarbons from carbon oxides and hydrogen, the catalyst comprising:

an acidic substrate comprising a zeolite, or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite or M-zeolite catalyst is 100 or more.

According to the invention, there is also provided a method or methods of preparing any of the catalysts described herein.

In some examples, the method includes producing a mixed catalyst, the method further including the step of mixing the dehydration/hydrogenation catalyst and a carbon oxide(s) conversion catalyst, for example a methanol synthesis catalyst. Preferably the mixed catalyst is adapted for the conversion of carbon oxide(s) and hydrogen to form saturated hydrocarbons, in particular C₃ and higher saturated hydrocarbons.

Thus the invention further provides the use of a catalyst as described herein in the catalysed conversion of carbon oxide(s) and hydrogen to form saturated hydrocarbons. According to a further aspect of the invention there is provided a process for the catalysed production of saturated hydrocarbons using a dehydration/hydrogenation catalyst comprising:

a zeolite or M-zeolite composition, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite is 100 or more.

Preferably the dehydration/hydrogenation catalyst is exposed to a source of a gas including methanol and/or DME and hydrogen. The catalyst may comprise the

dehydration/hydrogenation catalyst and a further catalyst, for example a carbon oxide(s) conversion catalyst, for example a methanol synthesis catalyst.

The reactants may for example comprise syngas. Preferably the process includes feeding syngas to the dehydration/hydrogenation catalyst.

The process is preferably in gas phase. The reaction temperature may be between from about 260 to 400 degrees C, for example from about 290 to 335 degrees C. The reaction pressure may be between from about 0.5 to 6.0MPa, for examples from 2.0 to 3.0MPa. The gas space velocity may be from about 500 to 6000h^_1, and for example about 1000 to 1500h^'\ Preferably the gas space velocity is defined as the hourly volume of gas flow in standard units divided by the catalyst volume.

In some examples, the carbon oxide(s) conversion catalyst may be in a first stage with a second stage including the dehydration/hydrogenation catalyst. The first and second stages may be physically separate, may be spaced apart from each other, or may have a direct interface, or may be spaced using a physical spacer, or other method. In such arrangements, the catalyst may be non-homogeneous in that there are carbon oxide(s) conversion catalyst rich regions (or region) and dehydration/ hydrogenation catalyst rich regions (or region). The carbon oxide(s) conversion catalyst region(s) will normally be arranged upstream of the dehydration/hydrogenation catalyst region(s). In examples, the process may include an upstream catalyst bed including the carbon oxide(s) conversion catalyst, for example for the production of DME and/or methanol from carbon oxides and hydrogen. Thus the process may be carried out in a multiple stage system. For example a carbon oxide(s) conversion catalyst, for example a methanol synthesis catalyst and/or DME synthesis catalyst may be provided in a first stage and the hydrogenation catalyst in a second stage. In some examples, the two stages will be separated. By separating the stages of the reaction system, it is possible to independently optimize the two stages. A significant advantage of this for some examples is that the methanol- and/or DME-generating catalyst can be run at conditions more suitable for improved conversion, selectivity, and/or longer catalyst life.

Also provided by an aspect of the invention is an integrated process for the generation of saturated C₃ and higher hydrocarbons from carbon oxide(s) and hydrogen, the process comprising the steps of: (a) feeding a gas stream including carbon oxide(s) and hydrogen to a reaction system comprising a catalyst system including a first stage catalyst composition, and a second stage catalyst composition, and (b) removing a product stream from the reaction system, the product stream including saturated C₃ and higher

hydrocarbons, wherein the first stage catalyst composition comprises a dimethyl ether (DME) synthesis catalyst and/or a methanol synthesis catalyst, and wherein the second stage catalyst composition comprises a zeolite or M-zeolite composition, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite is 100 or more.

The apparatus may comprise a one-stage reaction system and therefore the method may include the step of feeding the gas to a one-bed system comprising the first stage catalyst composition and the second stage catalyst composition. For example the one bed may include a mixture of the carbon oxide(s) conversion catalyst and the

dehydration/hydrogenation catalyst.

Preferably the temperature of the bed is more than 300 degrees C.

The carbon oxide(s) conversion catalyst may be active to produce methanol and/or active to produce DME.

The dehydration/hydrogenation catalyst may include a source of Pd and/or Cu. The dehydration/hydrogenation catalyst may include ZSM-5.

The method may include feeding the gas stream to a two-bed reaction system comprising a first bed including the first stage catalyst composition wherein the gas stream is at least partly converted in the first bed to form an intermediate product stream, and feeding at least a part of the intermediate product stream to a second bed including the second stage catalyst composition.

Preferably the first reaction stage temperature is lower than the second stage temperature, for example at least 20 degrees or at least 50 degrees lower. The temperature of the first stage may be less than 300 degrees C. Preferably, the temperature of the first stage is less than 295 degrees C, for example not more than 280 degrees C, for example not more than 250 degrees C. In examples, the temperature of the first stage may be between from about 190 to 250 degrees C, for example between from about 210 to 230 degrees C. In practical systems, it is likely that the temperature will vary across the reaction stage. Preferably the temperature of the stage is measured as an average temperature across a reaction region.

The temperature of the second stage may be more than 300 degrees C.

In some examples, the temperature of the second stage will be 320 degrees C or more. In some examples, a temperature of 340 degrees C or more will be preferred. In some examples the temperature of the second stage will be between from about 330 to 360 degrees C. In many cases it will be preferable for the temperature of the second stage to be less than 450 degrees C, for example less than 420 degrees C, or for example less than 400 degrees C which may prolong the life of the catalyst. Depending on the target products, other temperatures may be used for the second stage.

The first and second stages may be operated at the same or at different pressures. Both stages may be operated for example at a pressure less than 40 bar. In some examples, it will be preferable for the second stage to be operated at a pressure lower than that of the first stage, for example at least 5 bar lower, for example at least 10 bar lower.

For example, the first stage may be operated at a pressure of less than 40 bar, less than 20 bar, or less than 10 bar. In some examples, a significantly higher pressure may be desirable.

For example, the second stage may be operated at a pressure of less than 20 bar, less than 10 bar, or less than 5 bar. In some examples, a significantly higher pressure may be desirable.

The gas hourly space velocity of the first stage may be for example between about 500 and 6000, for example between about 500 and 3000.

The gas hourly space velocity of the second stage may be for example between about 500 and 20000, for example between about 1000- 10000.

Preferably the gas hourly space velocity is defined as the number of bed volumes of gas passing over the catalyst bed per hour at standard temperature and pressure.

Several configurations of the two stages are possible. An example giving less flexibility is one in which the two stages are contained within a single reactor vessel, for example as separate zones. In such a system, a heat transfer region may be provided, for example to control the reaction stage temperatures independently.

A more flexible system provides the two stages in separate vessels. At least a portion of the intermediate product stream (or effluent) exiting the first stage preferably passes directly to the second stage. In some examples, substantially the entire intermediate product stream passes to the second stage.

It will be understood that additional second stage influent components can be added to the intermediate stream upstream of the second stage. For example, addition of hydrogen and/or DME may be carried out. The intermediate stream may be subject to operations for example heat exchange upstream of the second stage and/or pressure adjustment, for example pressure reduction.

Each of the stages may include any appropriate catalyst bed type, for example fixed bed, fluidized bed, moving bed. The bed type of the first and second stages may be the same or different.

Potential application for example for the second stage is the use of a moving bed or paired bed system, for example a swing bed system, in particular where catalyst regeneration is desirable.

The feed to the process includes carbon oxide(s) and hydrogen. Any appropriate source of carbon oxides (for example carbon monoxide and/or carbon dioxide) and of hydrogen may be used. Processes for producing mixtures of carbon oxide(s) and hydrogen are well known. Each method has its advantages and disadvantages, and the choice of using a particular reforming process over another is normally governed by economic and available feed stream considerations, as well as by the desire to obtain the desired (¾- C0₂):(CO+C0₂) molar ratio in the resulting gas mixture, that is suitable for further processing. Synthesis gas as used herein preferably refers to mixtures containing carbon dioxide and/or carbon monoxide with hydrogen. Synthesis gas may for example be a combination of hydrogen and carbon oxides produced in a synthesis gas plant from a carbon source such as natural gas, petroleum liquids, biomass and carbonaceous materials including coal, recycled plastics, municipal wastes, or any organic material. The synthesis gas may be prepared using any appropriate process for example partial oxidation of hydrocarbons (POX), steam reforming (SR), advanced gas heated reforming (AGHR), microchannel reforming (as described in, for example, US Patent No. 6,284,217), plasma reforming, autothermal reforming (ATR) and any combination thereof.

A discussion of these synthesis gas production technologies is provided for in "Hydrocarbon Processing" V78, N.4, 87-90, 92-93 (April 1999) and/or "Petrole et Techniques", N. 415, 86-93 (July- August 1998), which are both hereby incorporated by reference.

The synthesis gas source used in the present invention preferably contains a molar ratio of (H₂-C0₂): (CO+C0₂) ranging from 0.6 to 2.5. The gas composition which the catalyst is exposed to will generally differ from such a range due to for example gas recycling occurring within the reaction system. For example, in commercial methanol plants, a syngas feed molar ratio (as defined above) of 2:1 is commonly used, whereas the catalyst may experience a molar ratio of greater than 5:1 due to recycle. The gas composition experienced by the catalyst in the first stage where a two-stage process is used may initially be for example between from about 0.8 to 7, for example from about 2 to 3.

Carbon oxide(s) conversion catalysts for example methanol synthesis catalysts are commonly water gas shift active. The water gas shift reaction is the equilibrium of H₂ and C0₂ with CO and H₂0. The reaction conditions for the methanol synthesis catalyst (for example in the first stage) preferably favour the formation of H₂ and C0_2. For the case where the carbon oxide(s) conversion catalyst is active to produce methanol, the reaction stoichiometry requires a synthesis gas molar ratio of 2:1. For the case where the carbon oxides(s) conversion catalyst is active to produce dimethyl ether (DME), the reaction coproduces water which is shifted with CO according to the water gas shift reaction to C0₂ and hydrogen. In case, the synthesis gas molar ratio (as defined above) requirement is also 2:1 but here a reaction product is C0₂. The second part of the reaction, for example the second stage reaction in the case of methanol synthesis in the first stage is thought to comprise initial conversion to DME and water, and subsequent conversion of DME to C₃ and higher saturated hydrocarbons and water. The second stage reaction in the case of DME synthesis in the first stage is thought to comprise only the stages of DME conversion to C₃ and higher saturated hydrocarbons and water. In this case, the product mixture additionally includes carbon dioxide. Where a hybrid catalyst is used, these two stages will be in the same reactor.

The catalysts of the methods may comprise any of the catalyst compositions described herein as appropriate.

The carbon oxide(s) conversion catalyst may be active to produce methanol in the first bed and or dimethyl ether (DME) in the first bed.

The temperature of the second bed may be more than 300 degrees C.

The carbon oxide(s) conversion catalyst may comprise a copper oxide and/or may include a zeolite and/or γ-Α1₂0₃.

The dehydration/hydrogenation catalyst may include a source of Pd and/or Cu. The second bed may include ZSM-5.

A further aspect of the invention provides an apparatus for carrying out a method as defined herein.

Also provided by the invention is apparatus for use in a process as described herein and a dehydration/hydrogenation catalyst obtained or obtainable by a method described herein.

The invention extends to methods and/or apparatus and/or catalyst composition substantially as herein described with reference to the accompanying drawings.

Any feature in one aspect of the invention may be applied to other aspects of the invention, in any appropriate combination. In particular, features of method aspects may be applied to apparatus aspects, and vice versa.

Preferred features of the present invention will now be described, purely by way of example, with reference to the accompanying drawing, in which:

Figure 1 shows schematically an example of a two-stage reactor system used in a process for the conversion of syngas to saturated hydrocarbons in an example of the invention.

The following describes example catalyst systems and example methods for their preparation and describes their evaluation in a single or two-stage reactor system. Figure 1 shows schematically an example of a two-stage test reactor system 1 for saturated hydrocarbon synthesis from syngas. The system 1 includes two reaction stages 3, 5 arranged in series. Each reaction stage 3, 5 includes a reaction vessel containing a fixed bed catalyst system. The reactions were carried out under pressurized conditions in these examples. Each stage 3, 5 was equipped with an electronic temperature controller for a furnace, a tubular reactor with an inner diameter of 12mm, and a back pressure valve 21, 21 ' downstream of the reactor. When carrying out examples including a one catalyst stage, only the reactor of the first stage 3 was used.

The upstream reaction stage 3 includes a first catalyst composition including a methanol synthesis catalyst; the downstream reactor vessel 5 contains a second catalyst composition including a dehydration/hydrogenation catalyst.

A syngas feed line 7 feeds syngas via a first pressure test point PI, a pressure reducing valve 9, a second pressure test point P2, a globe valve system including a mass flowmeter 11, and a third pressure test point P3 to the first reaction stage 3. A nitrogen feed line 13 is provided for feeding N₂ to a point at the first pressure test point PI . A hydrogen feed line 15 and vent 17 is provided upstream of the pressure reducing valve 9. Intermediate product stream leaving the first reaction stage 3 via line 19 passes through a back pressure valve to a fourth pressure test point P4 before passing to the second reaction stage 5. A product stream passes from the second reaction stage 5 via line 23 through a further back pressure valve 2 .

The system further includes gas chromatography (GC) apparatus 25 arranged to receive intermediate product stream from line 19 and/or product stream from line 23. The gas chromatography apparatus 25 in this example includes a flame ionization detector (FID) and a thermal conductivity detector (TCD).

Catalyst Evaluation

In the examples, the catalyst was first activated at 250 degrees C for 2 hours in a pure hydrogen flow. Subsequently, syngas was fed to the reaction vessels and the reaction was carried out using different reaction conditions as described below. All the products from the reactor were formed in the gaseous phase and analysed by gas chromatography on-line. The components CO, C0₂, CH₄ and N₂ were analysed using a GC equipped with a TCD and organic compounds were analyzed by another GC apparatus equipped with a FID. Experiment 1:

A commercial Cu-ZnO-Al₂0₃ (Cu-Zn-Al) methanol synthesis catalyst (from Shenyang Catalyst Corp.) was crushed into particles of size 20-40 mesh. ZSM-5 (from Nankai University Catalyst Ltd.) was pelletized and crushed into particles of size 20-40 mesh. 0.35g Cu-Zn-Al methanol synthesis catalyst and 0.45g ZSM-5 (SiO₂/Al₂O₃=140) were granular mixed. The catalytic performance was tested under 300 degrees C, 2.0MPa, 25ml/min syngas (H₂/CO=l/l).

CO conversion was 58.7%, selectivity of C0₂ was 52.5%, the sum selectivity of methanol and DME was 0.1% and the selectivity of hydrocarbons was 47.4%. The hydrocarbon distribution is shown in Table 1. The selectivity of C₅-Cn in gasoline was below 30% in this run, and the light gasoline C₅-C₈ product produced in this example had a relatively high content of iso-paraffms and relatively low content of aromatics and olefins. The aromatics C9-C11 also mainly contained trimethylbenzenes and tetramethylbenzenes. Table 1 Hydrocarbon distribution of Experiment 1

The figures in this table relate to the % selectivity of the products of the process, here measured as % C.

Experiment 2:

A commercial Cu-ZnO-A^Cb (Cu-Zn-Al) methanol synthesis catalyst was crushed into particles of size 20-40 mesh. ZSM-5 was pelletized and crushed into particles of size 20-40 mesh. 0.35g Cu-Zn-Al methanol synthesis catalyst and 0.45g ZSM-5

(SiO₂/Al₂O₃=140) were loaded in series into a reactor vessel to give an interface between the two components. The catalytic performance was tested under 290 degrees C, 2.0MPa, 25ml/min syngas (H₂/CO=l/l).

CO conversion was 32.9%, the selectivity of C0₂ was 51.5%, the sum selectivity of methanol and DME was 0.2% and the selectivity of hydrocarbons was 48.3%. The two components were loaded in series. Without wishing to be bound by particular theory, it is thought that because the components were loaded in series, the synergetic effect between them was weak and thus CO conversion was lower than that of Experiment 1. The selectivity of Cs-Cn increased to some extent. The hydrocarbon distribution is shown in Table 2.

Table 2 Hydrocarbon distribution of Experiment 2

Experiment 3:

A commercial Cu-ZnO-Al₂0₃ (Cu-Zn-Al) methanol synthesis catalyst and γ-Α1₂0₃ were powder mixed at the weight of 9/5, pelletized and crushed into particles of size 20-40 mesh. This catalyst was denoted as hybrid catalyst I. Pd modified ZSM-5 was prepared by the following ion-exchange method. lOg ZSM-5 was added to a 200ml solution of PdCl₂ at 60 degrees C with stirring, maintained for 8h, washed with water, dried at 120 degrees C, calcined at 550 degrees C and then pelletized and crushed into particles of size 20-40 mesh. The catalyst was denoted as 0.5%Pd-ZSM-5.

0.6g of hybrid catalyst I and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=140) was granular mixed. The catalytic performance was tested under 300 degrees C, 2.0MPa, 25ml/min syngas (H₂/CO=l/l). CO conversion was 72.3%. The selectivity of C0₂ was 52.8%. The sum selectivity of methanol and DME was 0.2%. The selectivity of hydrocarbons was 47.0%. The hydrocarbon distribution is shown in Table 3.

Table 3 Hydrocarbon distribution of Experiment 3

C₂-

Ci C₅-C₈ C₉-Cn

c₄

2.3 71.2 22.8 3.7

n- iso- C₉ C₁₀

olefin aromatic naphthene others alkane paraffin aromatic aromatic

8.0 82.2 0.0 2.5 7.3 31.5 63.1 5.4 Experiment 4:

0.6g of hybrid catalyst I and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=140) was loaded in series and there is an interface between the two components. The catalytic performance was tested under 300 degrees C, 3.0MPa, 35ml/min syngas (H₂/CO=l/l).

CO conversion was 62.3%. The selectivity of C0₂ was 50.1%. The selectivity of hydrocarbons was 49.9%. The hydrocarbon distribution is shown in Table 4. γ-Α1₂0₃ in hybrid catalyst I converted methanol formed on methanol synthesis catalyst into DME, so CO conversion was higher than that of Experiment 2. There was no substantial change of hydrocarbon distribution compared to Experiment 2.

Table 4 Hydrocarbon distribution of Experiment 4

Experiment 5:

0.6g of hybrid catalyst I was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=140) was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C, and the reaction temperature for the second catalyst bed was 320 degrees C. The reaction pressure was 3.0MPa, and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO conversion was 62.5%. The selectivity of C0₂ was 34.6%. The selectivity of hydrocarbons was 65.1%. The sum selectivity of methanol and DME was 0.3%. The hydrocarbon distribution is shown in Table 5.

Table 5 Hydrocarbon distribution of Experiment 5

C₂-

Ci C₅-C₈ C9-C₁₁

C₄

0.9 32.1 52.0 15.0

n- iso- C₉

olefin aromatic naphthene others alkane paraffin aromatic aromatic

4.0 74.6 2.7 10.1 8.6 42.7 48.1 9.2 Experiment 6:

0.6g of hybrid catalyst I was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=140) was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 294 degrees C, and the reaction temperature for the second catalyst bed was 330 degrees C. The reaction pressure was 4.0MPa, and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO conversion was 67.4%. The selectivity of C0₂ was 34.4%. The selectivity of hydrocarbons was 63.9%. The sum selectivity of methanol and DME was 1.7%. The hydrocarbon distribution is shown in Table 6.

Table 6 Hydrocarbon distribution of Experiment 6

Experiment 7:

A commercial Cu-ZnO-Al₂0₃ (Cu-Zn-Al) methanol synthesis catalyst and ZSM-5 (SiO₂/Al₂O₃=50) were powder mixed at the weight of 9/5, pelletized and crushed into particles of size 20-40 mesh. This catalyst is denoted as hybrid catalyst II.

0.6g of hybrid catalyst II was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5

(SiO₂/Al₂O₃=140) was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 265 degrees C and the reaction temperature for the second catalyst bed was 301 degrees C. The reaction pressure was 2.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l).

CO conversion was 47.3%. The selectivity of C0₂ was 38.1%. The selectivity of hydrocarbons was 61.8%. The sum selectivity of methanol and DME was 0.1%. The hydrocarbon distribution is shown in Table 7. The selectivity of C₅-Cn in the

hydrocarbons was higher when the two components were put separately in two catalysts beds. Table 7 Hydrocarbon distribution of Experiment 7

Experiment 8:

0.6g of hybrid catalyst II was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O =140) was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C and the reaction temperature for the second catalyst bed was 321 degrees C. The reaction pressure was 3.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO

conversion was 66.6%. The selectivity of C0₂ was 43.1%. The selectivity of hydrocarbons was 56.9%. There was no methanol and DME. The hydrocarbon distribution is shown in Table 8.

Table 8 Hydrocarbon distribution of Experiment 8

Experiment 9:

A commercial Cu-ZnO-Al₂0₃ (Cu-Zn-Al) methanol synthesis catalyst and ZSM-5 (SiO₂/Al₂O₃=140) were powder mixed at the weight of 9/5, pelletized and crushed into particles of size 20-40 mesh. This catalyst is denoted as hybrid catalyst III.

0.6g of hybrid catalyst III was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5

(SiO₂/Al₂O₃=140) was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C and the reaction temperature for the second catalyst bed was 321 degrees C. The reaction pressure was 3.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO

conversion was 50.2%. The selectivity of C0₂ was 35.9%. The selectivity of hydrocarbons was 64.1%. There was no selectivity of methanol and DME. The hydrocarbon distribution is shown in Table 9.

Table 9 Hydrocarbon distribution of Experiment 9

Experiment 10:

A commercial Cu-ZnO-Al₂0₃ (Cu-Zn-Al) methanol synthesis catalyst, γ-Α1₂0₃ and

ZSM-5 (SiO₂/Al₂O₃=140) were powder mixed at the weight of 9/5/10, pelletized and crushed into particles of size 20-40 mesh. This catalyst is denoted as hybrid catalyst IV. 1.65g of hybrid catalyst IV and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=140) was loaded in series and there was an interface between the two components. The catalytic performance was tested under 300 degrees C, 2.0MPa, 50ml/min syngas (H₂/CO=l/l).

CO conversion was 47.9%. The selectivity of C0₂ was 51.8%. The selectivity of hydrocarbons was 47.4%. The sum selectivity of methanol and DME was 0.8%. The hydrocarbon distribution is shown in Table 10.

Table 10 Hydrocarbon distribution of Experiment 10

Experiment 11:

0.6g of hybrid catalyst I, the granular mixture of hybrid catalyst I and 0.45g 0.5%Pd- ZSM-5 (SiO₂/Al₂O₃=140), and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=140) was loaded in series and there was an interface between the two components. The catalytic performance was tested under 290 degrees C, 2.0MPa, 50ml/min syngas (H₂/CO=l/l). CO conversion was 57.5%. The selectivity of C0₂ was 48.2%. The sum selectivity of methanol and DME was 0.4%. The selectivity of hydrocarbons was 51.4%. The hydrocarbon distribution is shown in Table 11.

Table 11 Hydrocarbon distribution of Experiment 11

Experiment 12:

Cu modified ZSM-5 was prepared by an incipient- wetness impregnation method. The Cu(N0₃)₂3H₂0 as the precursors of Cu was dissolved in water. About a 10ml solution was added to lOg ZSM-5 zeolite drop by drop in 5min, maintained for 24 h at room temperature, and then dried at 90 degrees C and calcined at 400 degrees C for 4 h. The resulting product is denoted herein as xCu-ZSM-5, where x stands for the weight content ofCu.

0.6g of hybrid catalyst I was in the first catalyst bed and 0.45g 5%Cu-ZSM-5 (SiO₂/Al₂O₃=140) was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C and the reaction temperature for the second catalyst bed was 320 degrees C. The reaction pressure was 3.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO

conversion was 60.9%. The selectivity of C0₂ was 35.7%». The selectivity of hydrocarbons was 60.9%). The sum selectivity of methanol and DME was 3.4%>. The hydrocarbon distribution is shown in Table 12.

Table 12 Hydrocarbon distribution of Experiment 12

C₂-

Ci C₅-C₈ C₉-Cn

C₄

1.1 43.4 46.7 8.8

n- iso- C₉ Cio

olefin aromatic naphthene others alkane paraffin aromatic aromatic

2.7 83.2 1.6 7.0 5.5 42.1 45.3 12.6 Experiment 13:

0.6g of hybrid catalyst I was in the first catalyst bed and 0.45g 10%Cu-ZSM-5 (SiO₂/Ai2O₃=140) was in the second catalyst bed. There is no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C and the reaction temperature for the second catalyst bed was 320 degrees C. The reaction pressure was 3.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO

conversion was 62.4%. The selectivity of C0₂ was 36.2%. The selectivity of hydrocarbons was 60.5%. The sum selectivity of methanol and DME was 3.3%. The hydrocarbon distribution is shown in Table 13.

Table 13 Hydrocarbon distribution of Experiment 13

Experiment 14:

0.6g of hybrid catalyst I was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=360) was in the second catalyst bed. There is no direct contact between the two components. The reaction temperature for the first catalyst bed was 294 degrees C and the reaction temperature for the second catalyst bed was 330 degrees C. The reaction pressure was 4.0MPa and the flow rate was 50ml/min syngas (H₂/CO=4/l).

CO conversion was 78.8%. The water gas shift reaction was restrained to some extent due to the high ratio of H₂/CO. The selectivity of C0₂ was only 23.7%. The selectivity of hydrocarbons was 76.3%. There is no methanol and DME. The hydrocarbon distribution is shown in Table 14.

Table 14 Hydrocarbon distribution of Experiment 14

C₂-

Ci C₅-C₈ C₉-Cn

C₄

1.1 43.7 41.8 13.4

n- iso- C₉ C₁₀

olefin aromatic naphthene others alkane paraffin aromatic aromatic

2.8 77.5 0.6 8.8 10.3 41.5 51.1 7.4 Experiment 15:

0.6 of hybrid catalyst I was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5

was in the second catalyst bed. There was no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C and the reaction temperature for the second catalyst bed was 320 degrees C. The reaction pressure was 3.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). A short term life test was carried out under such reaction conditions.

At time on stream of 1 h, CO conversion was 61.8%. The selectivity of C0₂ was 34.5%. The selectivity of hydrocarbons was 65.4%. The selectivity of methanol and DME was as low as 0.1%. The hydrocarbon distribution is shown in Table 15-1.

At time on stream of 9 h, CO conversion was 613%. The selectivity of C0₂ was 35.0%. The selectivity of hydrocarbons was 62.3%. The selectivity of methanol and DME was 2.7%. The hydrocarbon distribution is shown in Table 15-2.

At time on stream of 24 h, CO conversion was 58.9%. The selectivity of C0₂ was 34.4%. The selectivity of hydrocarbons was 58.2%. The selectivity of methanol and DME was 7.4%). The hydrocarbon distribution is shown in Table 15-3.

The selectivity of methanol and DME increased with time on stream. It may have been caused by the decreased dehydration ability of the catalyst. The content of olefins in the hydrocarbons increased and it may have been the result of the decreased hydrogenation ability of catalyst.

After 24h stability test, the reaction pressure are changed to 4.0MPa, and the reaction temperature increased to 330 °C of the second catalyst bed, and at time on stream of 31 h, CO conversion was increased to 67.4%. The selectivity of C0₂ was 34.4%. The selectivity of hydrocarbons was 63.8%. The content of C_5-11 in hydrocarbon was 69.7%. The selectivity of methanol and DME was 1.8%. The hydrocarbon distribution is shown in Table 15-4. Table 15 Hydrocarbon distribution of Experiment 15

Comparative Experiment 16:

0.6g of hybrid catalyst I was in the first catalyst bed and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O₃=50) was in the second catalyst bed. There is no direct contact between the two components. The reaction temperature for the first catalyst bed was 282 degrees C and the reaction temperature for the second catalyst bed was 320 degrees C. The reaction pressure was 3.0MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO conversion was 61.5%. The selectivity of C0₂ was 35.6%. The selectivity of hydrocarbons was 62. %. The sum selectivity of methanol and DME was 1.6%. The hydrocarbon distribution is shown in Table 16.

Table 16 Hydrocarbon distribution of Experiment 16

Experiment 17:

0.6g of hybrid catalyst I was loaded in the first reactor and 0.45g 0.5%Pd-ZSM-5 (SiO₂/Al₂O3=140) was loaded in the second reactor. The reaction temperature for the first reactor was 282 degrees C and the reaction temperature for the second reactor was 320 degrees C. The reaction pressure was 3.0 MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO conversion was 61.3%. The selectivity of C0₂ was 34.6%. The selectivity of hydrocarbons was 63.4%. The sum selectivity of methanol and DME was 2.0%. The hydrocarbon distribution is shown in Table 17

Table 17 Hydrocarbon distribution of Experiment 17 i C₂-C₄ C₅-C₈ C₉-Cn .4 35.1 53.8 9.6 n- iso- naph- C₉ C₁₀

olefin aromatic others alkane paraffin thene aromatic aromatic

4.1 75.7 4.6 9.0 6.5 46.3 34.6 19.1 Experiment 18:

0.6g of hybrid catalyst I was loaded in a first reactor and 0.45g l%Zn-ZSM-5

(SiO₂/Al₂O₃=140) was loaded in a second reactor. The reaction temperature for the first reactor was 260 degrees C and the reaction temperature for the second reactor was 350 degrees C. The reaction pressure was 3.0 MPa and the flow rate was 25ml/min syngas (H₂/CO=l/l). CO conversion was 59.9%. The selectivity of C0₂ was 33.2%. The selectivity of hydrocarbons was 66.7%. The sum selectivity of methanol and DME was 0.1%. The hydrocarbon distribution is shown in Table X

Table 18 Hydrocarbon distribution of Experiment 18

Further Comparative Example:

The catalyst bed was 0.5g H-ZSM-5. The reaction temperature was 400 degrees C. The reaction pressure was atmospheric pressure and the flow rate of N₂ was 25ml/min. The conversion of methanol was 100%. The selectivity of C_2-4 olefins in e_f fraction was about 66%. There are also many olefins existing in C_5-8. From Table 16, we can see that hydrocarbon distribution in the process when methanol as feedstock and nitrogen as carrier was very different from that in the above process when syngas as feedstock.

Table 19 Hydrocarbon distribution of Compared example 1

Ci c₂- C₅-C₈ C₉-Cn

C₄

0.4 50.3 41.7 7.6

n- iso- C₉ C₁₀

olefin aromatic naphthene others alkane paraffin aromatic aromatic

2.7 31.0 18.9 38.5 8.9 82.3 17.7 0 It will be understood that the present invention has been described above purely by way of example, and modification of detail can be made within the scope of the invention.

Each feature disclosed in the description, and (where appropriate) the claims and drawings may be provided independently or in any appropriate combination.

Claims

1. A catalyst composition for use as a dehydration/hydrogenation catalyst in a multistage catalyst system for the catalyzed production of C₅₊ saturated hydrocarbons from carbon oxide(s) and hydrogen, the catalyst composition comprising:

an acidic substrate comprising a zeolite, or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite or the zeolite of the M-zeolite catalyst is 100 or more.

2. A catalyst composition according to claim 1 , wherein the Si0₂/Al₂0₃ molar ratio of the zeolite of a second stage catalyst composition is 140 or more.

3. A catalyst composition according to claim 1 or claim 2, wherein the zeolite catalyst comprises ZSM-5.

4. A catalyst composition according to any preceding claim, wherein M comprises a metal chosen from the group comprising Pd, Pt, Rh, Ru, Cu and Zn, and preferably M comprises Pd and/or Cu.

5. A multi-stage catalyst system for use in the catalyzed production of C₅₊ saturated hydrocarbons from carbon oxides and hydrogen, the catalyst system comprising a first stage comprising a carbon oxide(s) conversion catalyst, and a second stage comprising a dehydration/hydrogenation catalyst comprising:

an acidic substrate comprising a zeolite, or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂ Al₂0₃ molar ratio of the zeolite or the zeolite of the M-zeolite catalyst is 100 or more.

6. A mixed catalyst for the catalysed production of C₅₊ saturated hydrocarbons from carbon oxides and hydrogen, mixed catalyst including:

a carbon oxide(s) conversion catalyst, and

a dehydration/hydrogenation catalyst comprising an acidic substrate comprising a zeolite or an M-zeolite catalyst, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite or the zeolite of the M-zeolite catalyst is 100 or more.

7. A catalyst according to claim 5 or claim 6 wherein the carbon oxide(s) conversion catalyst comprises a methanol synthesis catalyst.

8. A catalyst according to any of claims 5 to 7, wherein the carbon oxide(s) conversion catalyst comprises DME synthesis catalyst.

9. A catalyst according to any of claims 5 to 8, wherein the carbon oxide(s) conversion catalyst comprises a copper oxide.

10. A catalyst according to any of claims 5 to 9, wherein the carbon oxide(s) conversion catalyst comprises a zeolite and/or γ-Α1₂0₃.

11. A catalyst according to any of claims 5 to 9 wherein the dehydration/hydrogenation catalyst includes Pd and/or Cu.

12. An integrated process for the generation of saturated C₅ and higher hydrocarbons from carbon oxide(s) and hydrogen, the process comprising the steps of:

(a) feeding a gas stream including carbon oxide(s) and hydrogen to a reaction system comprising a catalyst system including a first stage catalyst composition, and a second stage catalyst composition, and

(b) removing a product stream from the reaction system, the product stream including saturated C₅ and higher hydrocarbons,

wherein the first stage catalyst composition comprises a dimethyl ether (DME) synthesis catalyst and/or a methanol synthesis catalyst,

and wherein the second stage catalyst composition comprises a zeolite, or M-zeolite composition, where M comprises a metal, wherein the Si0₂/Al₂0₃ molar ratio of the zeolite is 100 or more.

13. A process according to claim 12 comprising feeding the gas to a one-bed system comprising the first stage catalyst composition and the second stage catalyst composition.

14. A process according to claim 13, wherein the temperature of the bed is more than 300 degrees C.

15. A process according to any of claims 12 to 14, wherein the carbon oxide(s) conversion catalyst is active to produce methanol.

16. A process according to any of claims 12 to 15, wherein the carbon oxide(s) conversion catalyst is active to produce DME.

17. A method according to claim 12, including feeding the gas stream to a two-bed reaction system comprising a first bed including the first stage catalyst composition wherein the gas stream is at least partly converted in the first bed to form an intermediate product stream, and feeding at least a part of the intermediate product stream to a second bed including the second stage catalyst composition.

18. A process according to claim 17, wherein the carbon oxide(s) conversion catalyst is active to produce methanol in the first bed.

19. A process according to claim 17 or 18, wherein the carbon oxide(s) conversion catalyst is active to produce dimethyl ether (DME) in the first bed.

20. A process according to any of claims 17 to 19, wherein the temperature of the second bed is more than 300 degrees C.

21. A process according to any of claims 12 to 20, wherein the

dehydration/hydrogenation catalyst includes Pd and/or Cu.

22. A process according to any of claims 12 to 21 , wherein the

dehydration/hydrogenation catalyst includes ZSM-5.