The present invention relates to an integrated method for the co-production of both olefins and Gas-to-Liquids (GTL) products, and, in particular, to a method where the production of olefins is by autothermal cracking.

, Numerous processes are known for the cracking of a hydrocarbon feedstock to produce olefins. One such process is autothermal cracking, in which a hydrocarbon feed is mixed with oxygen and passed over an autothermal cracking catalyst. Combustion is initiated on the catalyst surface and the heat required to raise the reactants to process temperature and to carry out the endothermic cracking process is generated in situ. The product stream from the autothermal cracking process typically produces a gaseous stream comprising one or more olefins, oxygenates, carbon dioxide, methane, hydrogen and carbon monoxide. Such a process is described, for example, in EP 332289B; EP-529793B; EP-A-0709446 and WO 00/14035.

The olefins produced may themselves be used as feedstocks for olefin derivative processes, such as polymerization processes to produce polyethylene, polypropylene and other polymers. It is further desired to utilize the other products to generate as much value as possible. US 6,555,721 describes a process in which the synthesis gas components (hydrogen and carbon monoxide) from an autothermal cracking process are subsequently separated from the olefinic products and either (i) converted to methanol, (ii) reacted in a water-gas shift reaction or (iii) converted to hydrocarbons. The conversion of synthesis gas to hydrocarbons is generally known as the Fischer-Tropsch (FT) process, and US 6,555,721 also discloses that the lower boiling point hydrocarbons produced from this process may be recycled to the autothermal cracking process.

WO 03/066551 describes a process for the production of a mono-olefin from a feedstock comprising a paraffinic hydrocarbon by partially combusting at a pressure of at least 15 barg a mixture of the hydrocarbon feed and a molecular oxygen-containing gas in contact with a catalyst capable of supporting combustion beyond the normal fuel rich limit of flammability where they are reacted to form a product comprising one or more mono- olefϊn(s), carbon monoxide and hydrogen, and subsequently separating the product into a stream comprising carbon monoxide and hydrogen and a stream comprising one or more

olefins and recovering the one or more olefm(s). The stream comprising carbon monoxide and hydrogen may be subsequently passed to a Fischer-Tropsch reactor containing a Fischer-Tropsch catalyst wherein at least part of said stream is converted to hydrocarbons. We have now found that autothermal cracking processes and Gas-to Liquids (GTL) processes can be further advantageously integrated by being located in sufficiently close proximity to enable sharing of one or more process facilities (services).

Thus, the present invention provides a method for the co-production of product streams comprising hydrocarbons boiling in the diesel range and olefins respectively, said method comprising: (a) providing a Gas-to-Liquids (GTL) process and an autothermal cracking (ATC) process, said GTL process comprising a reforming reactor and a Fischer-Tropsch (FT) reactor, (b) reforming a methane-containing feed stream in said reforming reactor to produce a first reactant stream comprising synthesis gas, (c) reacting the first reactant stream comprising synthesis gas in said FT reactor to produce a product stream comprising hydrocarbons,

(d) treating the product stream comprising hydrocarbons to produce a first product stream comprising ethane and propane, and a second product stream comprising hydrocarbons boiling in the diesel range, (e) feeding the first product stream comprising ethane and propane to said autothermal cracking process, wherein said first product stream reacts with a molecular oxygen- containing gas in the presence of a catalyst to produce a product stream comprising olefins and synthesis gas,

(f) treating the product stream comprising olefins and synthesis gas to separate a (g) second reactant stream comprising synthesis gas and a product stream comprising olefins, and (h) feeding the second reactant stream comprising synthesis gas to the reformer of step

(b) or to the Fischer-Tropsch reactor of step (c), wherein said autothermal cracking process shares one or more process facilities with the GTL process.

In the method of the present invention the autothermal cracking process and the GTL process are provided in sufficiently close proximity that the autothermal cracking process

can share one or more process facilities (services) with the GTL process. Suitable process facilities include power generation, control room and flare facilities.

In a preferred embodiment, the autothermal cracking process and GTL process may share one or more gas and/or steam turbines used for power generation. For example, the reaction products from an autothermal cracking reaction are quenched as they emerge from the reaction chamber to avoid further reactions taking place. The heat from the quenching is used to generate high-pressure steam, which is converted to electricity via a turbine and used to provide power for those parts of the overall process requiring it. Because an ATC process generates relatively little steam (compared to a steam cracker for example), a turbine dedicated to an ATC process is typically fairly small. In the method of the present invention, steam produced in the ATC process may be exported to a much larger steam turbine on a GTL process. The method of the present invention also allows two turbines to be replaced by a single turbine. Thus, duplication is avoided and enhanced energy integration can achieved. In a further preferred embodiment the steam from the ATC process may be superheated in the heat recovery section of the GTL process.

A preferred, shared, turbine is a combined cycle cogeneration unit. In a typical cogeneration unit, the gas turbine consists of compressor and expander sections. Combustion air enters the compressor and is then contacted with the fuel gas in a combustion chamber. The hot combustion gases flow through the expander to provide the energy for the compressor and to drive an associated electric generator. Energy in the hot exhaust gases from the gas turbine is recovered by generation of high-pressure steam. This steam is used to produce additional electricity through a steam turbine. The overall efficiency of this system in converting fuel energy into electrical energy is about 50-55%, approximately double that of a typical steam cycle.

Thus, in a preferred embodiment of the present invention, the autothermal cracking process and GTL process share power generation facilities, and, most preferably, the autothermal cracking process and GTL process share a number of other pieces of equipment in addition, including control room and flare facilities. The present invention takes advantage of the fact that relative to "conventional" cracking processes, the economics of ATC processes are relatively less sensitive to the scale of the process and, hence, ATC can be operated economically at relatively small

scale. In contrast, the economics of steam cracking, for example, are much more sensitive to scale, and steam cracking rapidly becomes uneconomic unless operated at relatively large scale. This advantage of ATC allows smaller ATC processes to be built and operated at locations where it would not be economic to build and operate steam crackers, for example, because of lack of the required amounts of hydrocarbon feedstock. Thus, a GTL process and an autothermal cracking process can be advantageously co-located in locations where it would not be economically practical for a GTL process to be co-located with a steam cracker.

In particular, the method of the present invention may be operated in a manner where essentially all, and preferably 100% of, the paraffinic hydrocarbon feedstock to the ATC process is derived from products from the GTL process.

FT and ATC processes, generally, are known to the person skilled in the art, as described in US 6,555,721 and WO 03/066551.

The reforming of the methane-containing feed stream in step (b) of the method of the present invention preferably comprises catalytic steam reforming of methane. Typically, a supported nickel catalyst is employed.

The methane-containing feed stream may, for example, be methane as such or be methane in the form of natural gas.

In a most preferred embodiment of the present invention, prior to reforming, the methane-containing feed stream may be pre-saturated with water by contacting the methane-containing feed stream with an aqueous effluent derived from the ATC process. Pre-saturation of reformer feed streams is known in the art, but the use of the ATC aqueous effluent for the pre-saturation has the twin advantages of (a) reducing the absolute quantity of ATC aqueous effluent and (b) providing at least partial treatment of the ATC aqueous effluent to remove contaminants therein.

Where necessary, the methane-containing feed stream may be passed to a suitable pre-reformer prior to the "main" reforming reactor.

The carbon monoxide to hydrogen ratio of the synthesis gas suitable for the F-T process may suitably be in the range from 2:1 to 1:6, preferably from 2:1 to 1:2. Whilst the ratio of carbon monoxide to hydrogen in the synthesis gas produced by the reforming processes may differ from these ranges, it may be altered appropriately by the addition of

either component, or may be adjusted by the water-gas shift (WGS) reaction. The water gas shift reaction may be represented as the equilibrium:

CO + H ₂ O = CO ₂ + H ₂ (I)

The WGS reaction is generally operated in the presence of a catalyst; typically an iron oxide catalyst may be employed, although other catalysts may equally be used. Temperatures typically in the range from 350 to 500°C may suitably be used.

In general, the Fischer-Tropsch process of step (c) of the method of the present invention produces hydrocarbons from C ₁ upwards, principally up to C ₆ O in range. In recent years attention has been directed to the Fischer-Tropsch process as one of the more attractive direct and environmentally acceptable routes to high quality transportation fuels from alternative energy sources such as coal and natural gas via intermediate formation (reforming) of synthesis gas (carbon monoxide and hydrogen). The catalyst for the Fischer-Tropsch reaction may suitably comprise at least one metal selected from cobalt, nickel, iron, molybdenum, tungsten, thorium, ruthenium, rhenium, and platinum. Of the aforesaid metals cobalt, nickel and iron are preferred. Generally, the metals may be used in combination with a support material. Suitable support materials include alumina, silica and carbon, and mixtures of two Or more thereof. The use of cobalt, for example, as a catalytically active metal in combination with a support is known from, for example EP-A- 127220; EP-A-142887; GB-A-2146350; GB-A-2130113; EP-A-0209980; EP-A-0261870 and GB-A-2125062.

Fischer Tropsch conditions are suitably a temperature in the range from 160 to 350 ⁰ C, preferably from 180 to 275°C, and a pressure in the range from 0 to 100 barg, preferably from 5 to 50 barg e.g. 15-40barg. The GHSV for continuous operation may suitably be in the range from 100 to 25000 h ^'1 . The Fischer-Tropsch process may be carried out batchwise or continuously, preferably continuously, in a fixed bed, fluidised bed or slurry phase reactor. The Fischer- Tropsch process may also produce a fuel gas. Any residual carbon dioxide in the fuel gas can be removed, for example by passing the gas, together with any excess hydrogen, through a methanation stage. Heat from the reaction is used to produce a high pressure steam, which may be used to generate power for other parts of the process in one or more turbines as previously described.

In step (d) of the method of the present invention the product stream comprising hydrocarbons from the F-T reactor is treated to produce a first product stream comprising ethane and propane, and a second product stream comprising hydrocarbons boiling in the diesel range. This may be by any suitable method, typically by cooling the product stream and separating the liquid phase comprising the hydrocarbons boiling in the diesel range from the gaseous phase comprising the ethane and propane.

In step (e) of the process of the present invention, the first product stream comprising ethane and propane is fed to an autothermal cracking process, wherein said first product stream reacts with a molecular oxygen-containing gas in the presence of a catalyst to produce a product stream comprising olefins and synthesis gas.

The autothermal cracking process of the process is generally as described, for example, in EP-332289B; EP-529793B; EP-A-0709446 and WO 00/14035.

Combustion of a portion of the hydrocarbon feed with the oxygen occurs on contact with the autothermal cracking catalyst generating heat. The heat generated drives the subsequent dehydrogenation of the ethane and propane to produce the product stream comprising olefins and synthesis gas.

The catalyst is a catalyst capable of supporting combustion beyond the fuel rich limit of flammability. The catalyst usually comprises a Group VIII metal as its catalytic component. Suitable Group VIII metals include platinum, palladium, ruthenium, rhodium, osmium and iridium. Rhodium, and more particularly, platinum and palladium are preferred. Typical Group VIII metal loadings range from 0.01 to lOOwt %, preferably, between 0.01 to 20 wt %, and more preferably, from 0.01 to 10 wt % based on the total dry weight of the catalyst.

Where a Group VIII catalyst is employed, it is preferably employed in combination with a catalyst promoter. The promoter may be a Group IIIA, IVA, and/or VA metal.

Alternatively, the promoter may be a transition metal; the transition metal promoter being a different metal to that which may be employed as the Group VIII transition metal catalytic component.

Preferred Group IIIA metals include Al, Ga, In and Tl. Of these, Ga and In are preferred. Preferred Group IVA metals include Ge, Sn and Pb. Of these, Ge and Sn are preferred. The preferred Group VA metal is Sb. The atomic ratio of Group VIII B metal to the Group IIIA, IVA or VA metal may be 1 : 0.1 - 50.0, preferably, 1: 0.1 -

12.0.

Suitable metals in the transition metal series include those metals in Group IB to VIII of the Periodic Table. In particular, transition metals selected from Groups IB, IIB, VIB, VIIB and VIII of the Periodic Table are preferred. Examples of such metals include Cr, Mo, W, Fe, Ru, Os, Co, Rh, Ir, Ni, Pt, Cu, Ag, Au, Zn, Cd and Hg. Preferred transition metal promoters are Mo, Rh, Ru, Ir, Pt, Cu and Zn. The atomic ratio of Group VIII metal to transition metal promoter may be 1: 0.1 - 50.0, preferably, 1:0.1 - 12.0.

Preferably, the catalyst comprises only one promoter; the promoter being selected from Group IHA, Group IVA, Group VB and the transition metal series. For example, the catalyst may comprise a metal selected from rhodium, platinum and palladium and a promoter selected from the group consisting of Ga, In, Sn, Ge, Ag, Au or Cu. Preferred examples of such catalysts include Pt/Ga, WLn, Pt/Sn, Pt/Ge, Pt/Cu, Pd/Sn, Pd/Ge, Pd/Cu and Rh/Sn. The Rh, Pt or Pd may comprise between 0.01 and 5.0 wt %, preferably, between 0.01 and 2.0 wt %, and more preferably, between 0.05 and 1.0 wt % of the total weight of the catalyst. The atomic ratio of Rh, Pt or Pd to the Group III A, IVA or transition metal promoter may be 1 : 0.1 - 50.0, preferably, 1: 0.1 - 12.0. For example, atomic ratios of Rh, Pt or Pd tb Sn may be 1: 0.1 to 50, preferably, 1: 0.1 - 12.0, more preferably, 1: 0.2 - 3.0 and most preferably, 1: 0.5 - 1.5. Atomic ratios of Pt or Pd to Ge, on the other hand, may be 1: 0.1 to 50, preferably, 1: 0.1 - 12.0, and more preferably, 1: 0.5 - 8.0. Atomic ratios of Pt or Pd to Cu may be 1 : 0.1 - 3.0, preferably, 1 : 0.2 - 2.0, and more preferably, 1: 0.5 - 1.5.

Alternatively, the promoter may comprise at least two metals selected from Group IIIA, Group IVA and the transition metal series. For example, where the catalyst comprises platinum, the platinum may be promoted with two metals from the transition metal series, for example, palladium and copper. Such Pt/Pd/Cu catalysts may comprise palladium in an amount of 0.01 to 5 wt %, preferably, 0.01 to 2 wt %, and more preferably, 0.01 to 1 wt % based on the total weight of the dry catalyst. The atomic ratio of Pt to Pd may be 1: 0.1 - 10.0, preferably, 1: 0.5 - 8.0, and more preferably, 1: 1.0 -5.0. The atomic ratio of platinum to copper is preferably 1: 0.1 - 3.0, more preferably, 1: 0.2 - 2.0, and most preferably, 1: 0.5 - 1.5.

Where the catalyst comprises platinum, it may alternatively be promoted with one transition metal, and another metal selected from Group IIIA or Group IVA of the periodic

table. In such catalysts, palladium may be present in an amount of 0.01 to 5 wt %, preferably, 0.01 to 2.0 wt %, and more preferably, 0.05 - 1.0 wt % based on the total weight of the catalyst. The atomic ratio of Pt to Pd may be 1: 0.1 - 10.0, preferably, 1: 0.5 - 8.0, and more preferably, 1: 1.0 -5.0. The atomic ratio of Pt to the Group IIIA or IVA metal may be 1: 0.1 -60, preferably, 1 : 0.1 -50.0. Preferably, the Group IIIA or IVA metal is Sn or Ge, most preferably, Sn.

For the avoidance of doubt, the Group VIII metal and promoter in the catalyst may be present in any form, for example, as a metal, or in the form of a metal compound, such as an oxide. The catalyst may be unsupported, such as in the form of a metal gauze, but is preferably supported. Any suitable support may be used such as ceramic or metal supports, but ceramic supports are generally preferred. Where ceramic supports are used, the composition of the ceramic support may be any oxide or combination of oxides that is stable at high temperatures of, for example, between 600 ⁰ C and 1200 ⁰ C. The support material preferably has a low thermal expansion co-efficient, and is resistant to phase separation at high temperatures.

Suitable ceramic supports include corderite, lithium aluminium silicate (LAS), alumina (0C-AI ₂ O ₃ ), yttria stabilised zirconia, alumina titanate, niascon, and calcium zirconyl phosphate. The ceramic supports may be wash-coated, for example, with γ-Al ₂ C"3. The catalyst may be prepared by any method known in the art. For example, gel methods and wet-impregnation techniques may be employed. Typically, the support is impregnated with one or more solutions comprising the metals, dried and then calcined in air. The support may be impregnated in one or more steps. Preferably, multiple impregnation steps are employed. The support is preferably dried and calcined between each impregnation, and then subjected to a final calcination, preferably, in air. The calcined support may then be reduced, for example, by heat treatment in a hydrogen atmosphere.

The first product stream comprising ethane and propane may be fed to the autothermal cracking process as is, or may be fed in admixture other hydrocarbons and optionally other materials, for example methane, nitrogen, carbon monoxide, carbon dioxide, steam or hydrogen. Hydrogen is an especially preferred co-feed. Hydrogen

combusts to generate heat and thereby reduces the amount of paraffinic hydrocarbon combustion required to generate the required heat for dehydrogenation.

As the molecular oxygen-containing gas there may suitably be used either oxygen or air. It is preferred to use oxygen, optionally diluted with an inert gas, for example nitrogen. It is preferred to pre-mix the oxygen-containing gas and the first product stream, and optionally any other hydrocarbon-containing feeds, prior to contact with the catalyst. The composition of the feed gas mixture (paraffinic hydrocarbon to oxygen ratio) is suitably from 5 to 13.5 times the stoichiometric ratio of hydrocarbon to oxygen-containing gas for complete combustion to carbon dioxide and water. The preferred composition is from 5 to 9 times the stoichiometric ratio of hydrocarbon to oxygen-containing gas. The autothermal cracking process may suitably be operated at a temperature (catalyst exit temperature) greater than 500 ⁰ C, for example greater than 650 ⁰ C, typically greater than 750 ⁰ C, and preferably greater than 800 ⁰ C. The upper temperature limit may suitably be up to 1200 ⁰ C, for example up to 1100 ⁰ C, preferably up to 1000 ⁰ C. Most preferably, the autothermal cracking step is carried out at a catalyst exit temperature in the range 85O ⁰ C to 1000 ⁰ C.

It is preferred, although not essential, to preheat the feedstock and the oxygen- containing gas, suitably to 200 to 500 ⁰ C, preferably 200-300 ⁰ C.

The autothermal cracking process may be operated at atmospheric or elevated pressure. Pressures of 1 to 40 bara may be suitable. Preferably the autothermal cracking process is operated at a pressure of greater than 15barg. The pressure is preferably less than 50 barg. More preferably the autothermal cracking process is operated at a pressure of between 15-40barg and advantageously between 20-30barg e.g. 25barg.

Preferably, the gaseous feedstock and the molecular oxygen-containing gas are fed to the autothermal cracking process in admixture under a Gas Hourly Space Velocity

(GHSV) of greater than 80,000 hr ^"1 in order to minimise the formation of carbon monoxide and carbon dioxide. Preferably, the GHSV exceeds 200,000 hr ^"1 , especially greater than 1,000,000 hr ^"1 . For the purposes of the present invention GHSV is defined as:- vol. of total feed at NTP/Time/( vol. of catalyst bed). The most preferred gas hourly space velocity is pressure dependent and greater

than 100,000 h ^"1 barg ^'1 . For example, at 20 barg pressure, the gas hourly space velocity is most preferably, greater than 2,000,000 IT ¹ . It will be understood, however, that the optimum gas hourly space velocity will depend upon the nature of the feed composition. The product stream from the autothermal cracking process is usually quenched as it emerges from the reaction chamber to avoid further reactions taking place and the temperature of the stream is reduced to a temperature between 750-600 ⁰ C. Typically, the temperature of the stream is reduced to a temperature between 750-600°C within less than lOOmilliseconds of formation, preferably within 50milliseconds of formation and most preferably within 20milliseconds of formation e.g. within lOmilliseconds of formation. The product stream may be passed to at least one heat exchanger wherein the stream is cooled to a temperature approaching the dew point of the stream and the heat from the heat exchanger may be used to generate high-pressure steam, which may be used as described previously to generate power in turbines shared with the GTL process.

In addition to olefins, the autothermal cracking reaction produces hydrogen, carbon monoxide, methane, and small amounts of acetylenes, aromatics and carbon dioxide. The carbon dioxide is usually removed first, typically using an amine-based absorption system such as MEA or TEA (or mixtures of both), or any other commercially available CO ₂ removal process.

In step (f) of the method of the present invention the product stream is typically treated to separate the second reactant stream comprising synthesis gas, from the olefins, in a refrigeration facility (cryogenic separation unit) to separate methane, hydrogen and carbon monoxide. As noted further below, in one embodiment this refrigeration facility may be shared with that on an LNG process. Where not already at high pressure, the product stream may be compressed to a pressure between 15 and 40 barg to facilitate the separation of the products.

The olefin(s) that are separated are treated and recovered, typically by separation to produce olefin products and paraffinic products. The paraffmic products may be recycled to the autothermal cracking step. In a preferred embodiment, the heavier components of the product stream may be combined with the FT reaction products. For further details of preferred methods of operation of an ATC process reference may be made to EP-B 1-0332289; EP-B 1-0529793; and EP-A-0709446.

The second reactant stream comprising synthesis gas (carbon monoxide and hydrogen) separated in step (f) generally will have a hydrogen/carbon monoxide molar ratio of less than 2:1, typically having a molar ratio of hydrogen to carbon monoxide in the range 1:1 to 1.5:1. If necessary (typically where the second reactant stream comprising synthesis gas is fed directly to the Fischer-Tropsch reactor of step (c)) the ratio of carbon monoxide to hydrogen in the synthesis gas in the second reactant stream comprising synthesis gas may be altered by the addition of either component, or may be adjusted by the water-gas shift (WGS) reaction, as previously described..

In the process of the present invention, it is possible to utilise a smaller reformer reactor in step (b) than in a non-integrated process, since at least a portion of the synthesis gas components for the Fischer-Tropsch reaction are provided from the ATC process rather than from reforming of methane (methane-containing feed stream). This can have

« significant cost advantages since the reforming section of a GTL plant usually accounts for over half the costs of the plant. In a further embodiment of the present invention, the ATC and GTL processes may also be co-located with a liquefied natural gas (LNG) process, and, preferably, may also share at least some facilities with the LNG process.

For example, the ATC process, GTL process and LNG process may all share power generation facilities, flare facilities and/or control room facilities. As a further example, the ATC process and the LNG process may share refrigeration facilities. Refrigeration is generally used on an ATC process to separate methane, hydrogen and carbon monoxide from olefins in the autothermal cracking product stream.

Refrigeration facilities are generally more economical the larger the scale of the facility.

Thus, by sharing refrigeration facilities between an ATC process and an LNG process two facilities may be replaced by a single facility. In addition, because a refrigeration facility for an LNG process is typically significantly larger than a facility for an autothermal cracking process, the removal of a separate refrigeration facility from an autothermal cracking process can be achieved without requiring a significantly larger refrigeration facility for the LNG process. The present invention is illustrated with respect to Figure 1, wherein Figure 1 is a schematic diagram of an integrated GTL and ATC process.

With reference to Figure 1, natural gas (1) and steam (2) are fed to a steam reforming reactor (3a) of a GTL process (3) to produce a stream comprising synthesis gas (4), which is then fed to a Fischer-Tropsch reactor (3b) of the (GTL) process (3), to give a product stream comprising hydrocarbons (5). This product stream is cooled in a suitable separator (6) to produce a product stream comprising ethane and propane (7) and a product stream comprising hydrocarbons boiling in the diesel range (8). The product stream comprising ethane and propane is passed to an autothermal cracking process (9) wherein it is reacted with an oxygen-containing gas (10) in the presence of a catalyst capable of supporting combustion beyond the normal fuel rich limit of flammability to produce a product stream comprising olefins and synthesis gas (11), which is passed to a cryogenic separation unit (12). Steam generated in the autothermal cracking process (9) is passed to a gas turbine on the GTL process (shown schematically by the line 13).

In the cryogenic separation unit (12), the product stream comprising olefins and synthesis gas (11) is treated to separate a product stream comprising synthesis gas (14) and a product stream comprising olefins (15). The product stream comprising synthesis gas (14) is passed to the F-T reactor (3b). In an alternative embodiment, represented by the line 14b, all or part of this stream may instead be passed to the reforming reactor (3a).