This application is the U.S. national phase of international application PCT/GB01/04461, filed 5 Oct. 2001, which designated the U.S. 
    
    
     The present invention relates to a process for isomerising hydrocarbons. In particular, the present invention relates to a process for selectively isomerising hydrocarbons to obtain triptane (2,2,3-trimethyl butane). 
     BACKGROUND OF THE INVENTION 
     Triptane is a highly branched hydrocarbon which can be used as a fuel additive for motor and aviation gasoline because of its high octane rating. It is generally produced by isomerising or reacting aliphatic, cycloaliphatic and/or alkyl aromatic hydrocarbons in the presence of an isomerisation catalyst. An example of such a process is described in U.S. Pat. No. 3,766,286. This document describes the use of reaction temperatures of −30 to 100° C. and contact times of 0.1 to 10 hours. The preferred embodiment, for instance, describes the isomerisation of n-heptane at 25° C. for 5 to 6 hours. The triptane yields obtained were between 0.4 and 1.4 wt %. 
     SUMMARY OF THE INVENTION 
     We have now found that by careful selection of reaction temperature and contact time, the selectivity of the isomerisation process towards triptane may be increased. 
     According to the present invention, there is provided a process for the production of triptane, said process comprising: 
     isomerising a hydrocarbon feedstock by contacting said feedstock with an isomerisation catalyst at a reaction temperature of −50 to 25° C., and a contact time of 0.01 to 150 hours, such that the triptane selectivity of the isomerisation reaction is at least 5% as a proportion of said hydrocarbon feedstock. 
     Preferably, the reaction temperature is −30 to 15° C., more preferably, −25 to 10° C., even more preferably, −15 to 5° C., most preferably, −10 to 0° C. 
     The contact time may be 0.05 to 50 hours, preferably, 0.08 to 24 hours, more preferably, 0.1 to 15 hours, even more preferably, 1 to 10 hours, yet more preferably, 2 to 7 hours and most preferably, 4 to 6 hours. 
     Preferably, the triptane selectivity is at least 7%, more preferably, at least 9%. For example, the triptane selectivity may be between 9 and 60% of the initial hydrocarbon feedstock. 
     The isomerisation catalyst employed is preferably a superacid. Suitable superacids include those comprising a Lewis acid of the formula MX n , where M is an element selected from the transition metal series, Group 13, 14, 15 or 16 of the Periodic Table, X is a halogen, or an anion derived from fluorosulfuric acid, trifluoromethanesulfonic acid, or trifluoroacetic acid. n is an integer of 3 to 6. Bronsted acids selected from HX, where X has the same meaning as discussed above and mixtures thereof may also be employed. 
     Preferably, M is selected from Groups 13 and 15 of the Periodic Table. More preferably, M is Sb. X may be F, Cl, Br or I, and is preferably F or Cl. In preferred embodiments of the invention, M is employed in its highest valency state with the selected halogen. For example, in a most preferred embodiment of the invention, the Lewis acid is SbF 5 . 
     M may alternatively be selected from the transition metal series. Where M is a transition metal, it is preferably a metal from Group IV or V of the transition metal series. Preferred transition metals include Ti, Zr, Hf, V, Nb, Ta. More preferably, M is selected from Ti, Nb and Ta, and most preferably, M is Ta. 
     Preferably, the Bronsted acid is selected from HF, trifluoromethanesulfonic and fluorosulfuric acid. 
     Preferred examples of suitable isomerisation catalysts are HSO 3 F—SbF 5  and SbF 5 —HF. 
     The molar ratio of Bronsted acid to Lewis acid can range from about 20:1 to 1:5. Preferably, a 5:1 to 1:1 molar ratio is employed. The amount of the catalyst employed with reference to the total amount of hydrocarbon used may range from about 0.01 to 100 parts by weight of the catalyst per part by weight of hydrocarbon. Preferably, the amount of catalyst employed is 1 to 10 parts by weight of the catalyst per part by weight of the hydrocarbon. 
     The catalyst may be used as the neat liquid, as a diluted solution or adsorbed on a solid support. With regard to the diluted catalyst, any diluent may be used that is inert under the reaction conditions. To obtain optimum results, the diluents may be pretreated to remove catalyst poisons such as water, unsaturated compounds and the like. Typical diluents include sulfuryl chloride fluoride, sulfuryl fluoride, fluorinated hydrocarbons and mixtures thereof. Protic acids including fluorosulfuric acid, surfuric acid, trifluoromethanesulfonic acid and the like, themselves, can be used as diluents. The diluent:catalyst volume ratio can range from about 50:1 to 1:1 and, preferably, from 10:1 to 2:1. 
     The catalyst may alternatively be incorporated with a suitable solid carrier or support. Any solid catalyst support may be used that is substantially inert to the catalyst under the reaction conditions. The support may be pretreated, such as by heating, chemical treatment or coating, to remove substantially all water and/or hydroxylic sites that might be present. Active supports may be rendered inert by coating them with an inert material such as antimony trifluoride or aluminium trifluoride. Suitable solid supports include carbon (eg graphite), fluoride-treated or coated resins such as sulfonated cation exchange resins, fluoride-treated acidic chalcides such as alumina and aluminosilicates, and acid-resistant molecular sieves such as a zeolite, e.g. faujasite. The supported catalysts can be prepared in any suitable manner, such as by conventional methods including dry mixing, coprecipitation or impregnation. In one embodiment, the supported catalyst is prepared by impregnating a suitable deactivated support with a metal fluoride such as antimony pentafluoride and then with a Bronsted acid such as fluorosulfuric acid. 
     When a supported catalyst is employed, the weight ratio of the Lewis acid to the support may range from 1:100 to 1:10 and preferably, from 1:50 to 1:35. The weight ratio of the Bronsted acid to the support may range from 1:100 to 1:10 and, preferably, from 1:50 to 1:35. 
     The hydrocarbon feedstock that may be used in the instant process include paraffins, alkyl substituted aromatic hydrocarbons and mixtures thereof. The paraffins as herein defined include aliphatic and cycloaliphatic hydrocarbons that are substantially in the liquid phase at room temperature. The aliphatic hydrocarbons (straight and branched chain materials) can contain 4 to 20 carbon atoms per molecule, preferably 4 to 8 carbon atoms, and may be exemplified by n-butane, n-pentane, methylpentane, methylhexane and the like. The cycloaliphatic hydrocarbons (naphthenes) can contain 6 to 20 carbon atoms per molecule, preferably 6-12 carbon atoms, and may be exemplified by methylcyclopentane, the dimethylcyclopentanes, ethyleyclohexane, n-pentylcyclohexane and the like. Depending on reaction conditions, summarization of the ring, i.e. ring expansion or contraction, may compete with side chain isomerisation. The alkyl substituted aromatic hydrocarbons can contain 7 to 20 carbon atoms per molecule, preferably 7 to 12 carbon atoms, and may include in principle all isomerisable alkylaromatic or polyalkylaromatic hydrocarbons such as the xylenes, n-butyl benzene and the like. Both positional isomerisation of the alkyl groups substituted on the ring and side chain isomerisation may occur depending on reaction conditions. Other aliphatic or alicyclic hydrocarbons commonly found in conventional petroleum hydrocarbon light naphtha streams may also be present. 
     In a preferred embodiment of the invention, the hydrocarbon feedstock comprises a C 5  to C 9  alkane, for example, a C 7  alkane. C 7  naphtha streams, for instance, may be employed. Examples of suitable C 7  alkanes include n-heptane, 2-methyl hexane, 3-methyl hexane, ethyl pentane, 2,3-dimethyl pentane, 3,3-dimethyl pentane. 2,2-diemethyl pentane and 2,4-dimethyl pentane. The hydrocarbon feedstock may consist essentially of one of these C 7  alkanes, or may comprise a mixture of two or more of them. Suitable alkane mixtures include mixtures of 3-methyl hexane and 2,3-dimethyl pentane and mixtures of heptane and 2,4-dimethyl pentane. The C 7  alkanes may be present in combination with other hydrocarbon species such as naphthenes (eg 0 to 40%, preferably 30 to 36%), and aromatics, such as toluene (eg 0 to 10%, preferably 2 to 5%). An example of such a combination comprises 25 to 40%, preferably, 32 to 38% n-heptane; 10 to 28%, preferably, 15 to 23% mono-branched heptane; 5 to 15%, preferably, 7 to 11% dibranched heptane; 20 to 40%, preferably, 32 to 38% naphthenes and 0 to 5%, for example, 2 to 3% aromatics. Preferably, however, the aromatic content of the feeds is low, for example, less than 1%. The C 7  alkanes may also be present in the feed together with other alkanes such as C 5 , C 6  and C 8  alkanes. An example of such a mixture is an alkylate-C 7  stream comprising 2 to 6, preferably, 4% i-C 5 ; 3 to 7, preferably, 5% C 6 ; 50 to 70, preferably, 60 to 62% 2,3-dimethyl pentane; 20 to 30, preferably, 22 to 26% 2,4-dimethyl pentane and 2 to 8, preferably, 4 to 6% C 8 . 
     The hydrocarbon feedstock may contain various cracking inhibitors or moderators such as hydrogen and/or iso-butane. The inhibitors act to depress excessive cleavage reactions that may occur during the isomerisation. When hydrogen or iso-butane is used, it is employed in amounts ranging, preferably, from 1 to 3 mole percent based on hydrocarbon feed. 
     The process of the invention may be conducted as a batch or continuous type operation. In general, the various means customarily employed in extraction processes to increase the contact area between the hydrocarbon phase and the catalyst phase may be used. In one embodiment of the invention, the hydrocarbon phase and catalyst phase may be contacted substantially in the liquid phase. The apparatus employed may be of a conventional nature. For example, the apparatus may comprise a single reactor, such as a fluidised-bed reactor, or multiple reactors provided, for example, with efficient stirring devices, such as mechanical agitators, ultra-sonic agitators, jets of restricted internal diameter, and turbo mixers. The hydrocarbon phase and the catalyst phase may be passed through one or more reactors in concurrent, cross-current or counter-current flow. Unreacted reactants, catalysts, inhibitors and heavier products of the reaction may be separated from the desired isomeric product and from one another such as by distillation and returned in whole or in part to the isomerisation reaction. The resultant product may be further processed as by alkylation and the like, or be employed directly as a high octane gasoline blending agent. 
     The triptane product may be recovered from the product mixture using any suitable technique. Examples include distillation, extractive distillation, and selective crystallisation. Membranes may also be employed. 
     According to a further aspect of the present invention, there is provided a continuous process for the production of triptane, said process comprising: 
     feeding a hydrocarbon feedstock into a reactor, 
     contacting the feedstock with an isomerisation catalyst, under reaction conditions effective to produce a product mixture comprising triptane, 
     characterised in that the product mixture comprises at least two liquid phases, one liquid phase being denser than the other. 
     The two liquid phases may be separated by simple separation techniques, such as decantation. This separation step may be carried out continuously or at periodic intervals. 
     The denser of the two liquid phases is typically a polar or ionic phase comprising the isomerisation catalyst, and optionally, any catalyst diluent and/or catalyst support employed in the reaction. The denser liquid phase may be retained in the reactor. Alternatively, the denser liquid phase may be recovered from the reactor, and recycled, preferably, once at least some of the catalyst diluent has been removed from the recovered phase. The catalyst may also be regenerated prior to recycling. It should be noted that the denser of the two phases may be present as an emulsion. 
     The less dense of the two liquid phases is typically a less or non-polar phase comprising the triptane product. Optionally, other oily products, such as by-products of the reaction may also be present in the second phase. Examples of possible by-products include aliphatic hydrocarbons, for example, those comprising 3 to 10 carbon atoms. Other by-products include aromatic and polymeric species (C 12  and above). Such species may be fluorinated and/or sulphonated. In a preferred embodiment, the less dense phase is separated from the denser phase, and recovered from the reactor. 
     In addition to the two liquid phases, the product mixture may also comprise a vapour phase. The vapour phase may comprise triptane, other light aliphatic and aromatic hydrocarbons (eg C 1  to C 9 ), hydrogen and iso-butane. In a preferred embodiment, at least some of the vapour phase is withdrawn from the reactor. The vapour phase may be purified by condensation and distillation to produce a triptane-containing stream. 
     Any triptane recovered from the vapour phase may be used in the production of motor or aviation gasoline, especially, unleaded motor or unleaded aviation gasoline. In a preferred embodiment, the condensed vapour is purified further, for example, by distillation, to enhance its concentration of triptane. The remainder of the condensed vapour may be recycled to the reactor. At least one motor or aviation gasoline additive may then be added to the triptane-enhanced product. The resulting mixture may be employed as, or as an additive for a motor or aviation gasoline, preferably, an unleaded motor or aviation gasoline. 
     The contents of the reactor may be mixed. This mixing step may be carried out using any suitable technique, for example, by using a mechanical stirrer, an ultrasonic agitator and/or by introducing a gas or liquid into the reactor. Any suitable mechanical stirrer may be employed. Gases that may be bubbled through the reactor to agitate its contents include nitrogen, argon, hydrogen and light hydrocarbons (eg methane, isobutane). Additionally or alternatively, the mixing may be achieved simply as a result of the reactants and/or catalyst being introduced into the reactor. 
     Although mixing is important for facilitating reaction, it can also inhibit the separation of the two liquid phases. This problem may be alleviated by reducing the rate of agitation. Preferably, however, at least a portion of the product mixture is at least partially shielded from the full force of the agitation, so that it can separate into at least two liquid phases. 
     Thus, according to a preferred embodiment, there is provided a continuous process for the production of triptane said process comprising: 
     providing a reactor having a reaction zone and a separation zone, 
     feeding a hydrocarbon feedstock into a reactor, 
     contacting the feedstock with an isomerisation catalyst, under reaction conditions effective to produce a product mixture comprising triptane, 
     said process being characterised by having at least a portion of the product mixture in the separation zone, so that it can separate into at least two liquid phases. 
     The reaction zone and separation zone are preferably in fluid communication with each other. The reaction zone and separation zone may be provided in a single piece of apparatus, for example, by using a reactor having a reaction zone and a separation zone. An advantage of this arrangement is that it keeps pipe-work requirements to a minimum, reducing the material cost of the overall reactor. It should be noted, however, that it is possible to provide the reaction zone and separation zone using separate pieces of apparatus, for example, by coupling a reactor to a separation tank. Multiple reaction zones and/or separation zones may be employed. For example, a reactor having a reaction zone and a separation zone may be coupled to a separate separation tank. 
     In preferred embodiments of the invention, a reactor having at least one reaction zone and at least one separation zone is employed. For example, the reaction and separation zones may be separated using one or more grids and/or perforated plates. In use, the product mixture is allowed to flow freely between the reaction and separation zones through the apertures or perforations in the grid/plate. When the contents of the reactor on one side of the plate/grid is mixed, the reactor contents on the opposite side of the plate/grid is shielded at least in part from the fall force of the mixing. Thus, the reactor contents on the opposite side of the grid is in the separation zone, and can separate into at least two phases. The stirrer may be employed in combination with one or more baffles, which may be located in the reactor to enhance the mixing effect of in the reaction zone. 
     The grid or perforated plate may be located in the reactor and placed 0 to 60°, preferably, 0 to 45°, more preferably, 0 to 30° and most preferably, 0 to 15° to the horizontal. In one embodiment, the grid or plate is positioned substantially horizontally. The reactor contents below the grid or plate is agitated, allowing the separation zone to form above the grid or plate. Preferably, the edge(s) of the grid or plate is adjacent to the inner walls of the reactor. The edge(s) may be spaced or in physical engagement with the inner walls of the reactor. 
     In the embodiment described above, the contents of the reactor is allowed to flow relatively freely through the apertures or perforations of the grid/plate. In an alternative embodiment, the flow of product mixture from the reaction zone to the separation zone may be driven by a mechanical impellor or by the gas lift effect of any bubbles in the reactor. Flow through the separation zone in this mode may be controlled, for example, by positioning a barrier or weir between the reaction zone and separation zone and controlling the driving force across it by controlling the liquid/vapour or liquid/liquid interface levels on either side. Thus, product mixture is allowed to flow from the reaction zone to the separation zone either continuously or at periodic intervals. 
     The hydrocarbon feedstock may be introduced continuously or at periodic intervals. Preferably, the feedstock is fed to the reactor continuously. The feedstock may be fed into the reactor at a rate of more than 25 g of hydrocarbon per kg of catalyst per hour, for example, from 50 to 2500 g of hydrocarbon per kg of catalyst per hour. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
       These and other aspects of the present invention will now be described with reference to the drawings, in which: 
         FIG. 1  is a schematic diagram of an apparatus suitable for carrying out an embodiment of the process of the invention, and 
         FIG. 2  is a schematic diagram of an alternative apparatus suitable for carrying out an embodiment of the process of the invention. 
     
    
    
     DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS 
     Referring to  FIG. 1 , the apparatus comprises a reactor  10 , which is divided into a reaction zone  12  and a separation zone  14  by a grid  16 . The reactor is provided with a pair of baffle plates  18 ,  20  and a mechanical stirrer  22 . The stirrer  22  extends into the reaction zone  12 . 
     In operation, the reaction zone  12  is charged with 30 mol % SbF 5  in FSO 3 H via line “A”. A reactant stream comprising naphtha is also continuously fed to the reaction zone  12  via line “B”, and the contents of the reaction zone  12  is agitated by the mechanical stirrer  22 . The reaction zone  12  is maintained at −30 to 10° C., and a pressure of less than 50 bar. Under the reaction conditions, the naphtha is isomerised to produce a product mixture comprising triptane. 
     The contents of the reactor  10  is free to flow between the reaction zone  12  and the separation zone  14  through the apertures (not shown) of the grid  16 . The grid  16 , however, shields the product mixture in the separation zone at least in part from the full force of the agitation caused by the stirrer  22 . Thus, the product mixture in the separation zone  14  is allowed to settle and separate into a less dense phase, and a dense phase. The less dense phase contains the triptane product, and is continuously recovered from the separation zone via line “C”. The dense phase may be present as an emulsion because of the agitational forces in the reaction zone  12 . 
       FIG. 2  depicts an alternative apparatus for carrying out an embodiment of the present invention. The apparatus comprises a reactor  110 , which is divided into a reaction zone  112  and a separation zone  114  by a separator  116 . The separator  116  comprises an inlet  118 , and an outlet  120 . The reactor  110  is also provided with a mechanical stirrer  122 , which extends into the reaction zone  112 . 
     In operation, the reaction zone  112  is charged with 30 mol % SbF 5  in FSO 3 H via line “A”. A reactant stream comprising naphtha is also continuously fed to the reaction zone  112  via line “B”, and the contents of the reaction zone  112  is agitated by the mechanical stirrer  122 . The reaction zone  112  is maintained at −30 to 10° C., and a pressure of less than 50 bar. Under the reaction conditions, the naphtha is isomerised to produce a product mixture comprising triptane. 
     The reaction mixture of the reaction zone  112  is present as an emulsion because of the agitational forces caused by the stirrer  122 . The reaction mixture of the reaction zone  112 , however, is free to flow from the reaction zone  112  into the separation zone  114  via the outlet  120  of the separator  116 . In the separation zone, the reaction mixture is shielded from the full force of the agitation caused by the stirrer  122 . Thus, the reaction mixture is allowed to settle and separate into three phases: a less dense phase  124 , a middle phase  126  and a dense phase  128 . The less dense phase  124  contains the triptane product, and is continuously recovered from the separation zone via line “C”. The dense phase  128  comprises the acid catalyst (30 mol % SbF 5  in FSO 3 H). Some of this phase  128  is allowed to flow back into the reaction zone  112  via the inlet  118 . 
     EXAMPLES 
     A hydrocarbon (100 ml) was intimately mixed with a dense catalyst consisting of 30 mol % SbF 5  in FSO 3 H in a batch autoclave at temperature of −30 to 10° C. To ensure good mixing a baffle was employed, and the reaction mixture was stirred at 2000 rpm. Samples were removed from the reaction mixture at timed intervals and analysed by gas chromatography. The results are shown in Table 1 below. 
     
       
         
               
               
               
               
             
               
               
               
               
               
               
               
             
               
               
               
               
               
               
               
             
           
               
                   
                   
               
               
                   
                 Selectivity 
                   
                 Yield 
               
             
          
           
               
                   
                 Temp 
                 Res. time 
                 223TMB 
                 C 7  Isom 
                 Conv. 
                 223TMB 
               
               
                 Feed 
                 ° C. 
                 minutes 
                 % 
                 % 
                 % 
                 wt % 
               
               
                   
               
             
          
           
               
                 Heptane 
                 −30 
                 7440 
                 12.8 
                 94.4 
                 68.3 
                 8.7 
               
               
                 ″ 
                 −20 
                 1890 
                 11.4 
                 94.2 
                 45.3 
                 5.2 
               
               
                 ″ 
                 −10 
                 300 
                 10.4 
                 93.8 
                 50.5 
                 5.3 
               
               
                 ″ 
                 0 
                 120 
                 8.6 
                 91.4 
                 26.8 
                 2.3 
               
               
                 ″ 
                 10 
                 25 
                 8.5 
                 94.6 
                 9.9 
                 0.8 
               
               
                 3MHx 
                 −30 
                 4260 
                 15.9 
                 99.8 
                 87.0 
                 13.8 
               
               
                 ″ 
                 −20 
                 525 
                 15.1 
                 92.7 
                 85.1 
                 12.9 
               
               
                 ″ 
                 −10 
                 90 
                 10.0 
                 96.1 
                 82.1 
                 8.2 
               
               
                 ″ 
                 0 
                 36 
                 8.0 
                 97.7 
                 80.0 
                 6.4 
               
               
                 23DMP 
                 −30 
                 4260 
                 14.4 
                 99.8 
                 90.0 
                 13.0 
               
               
                 ″ 
                 −20 
                 1440 
                 13.9 
                 91.4 
                 92.2 
                 12.8 
               
               
                 ″ 
                 −10 
                 310 
                 13.4 
                 93.6 
                 92.1 
                 12.3 
               
               
                 ″ 
                 0 
                 75 
                 11.7 
                 91.6 
                 91.9 
                 10.8 
               
               
                 3MHx/ 
                 −30 
                 4860 
                 16.2 
                 99.7 
                 78.0 
                 12.6 
               
               
                 23DMP 
               
               
                 3MHx/ 
                 −20 
                 1020 
                 15.1 
                 99.7 
                 77.0 
                 11.6 
               
               
                 23DMP 
               
               
                 3MHx/ 
                 −10 
                 310 
                 14.0 
                 90.4 
                 72.7 
                 10.2 
               
               
                 23DMP 
               
               
                 3MHx/ 
                 0 
                 60 
                 10.0 
                 99.6 
                 76.0 
                 7.6 
               
               
                 23DMP 
               
               
                 Heptane/ 
                 −30 
                 4200 
                 14.5 
                 99.8 
                 59.0 
                 8.6 
               
               
                 24DMP 
               
               
                 Heptane/ 
                 −20 
                 1044 
                 13.8 
                 99.6 
                 58.0 
                 8.0 
               
               
                 24DMP 
               
               
                 Heptane/ 
                 −10 
                 240 
                 12.2 
                 99.6 
                 54.0 
                 6.6 
               
               
                 24DMP 
               
               
                 Heptane/ 
                 0 
                 60 
                 10.6 
                 99.5 
                 51.0 
                 5.4 
               
               
                 24DMP 
               
               
                 24DMP 
                 −30 
                 3600 
                 15.6 
                 99.8 
                 75.0 
                 11.7 
               
               
                 ″ 
                 −20 
                 900 
                 15.0 
                 99.7 
                 75.0 
                 11.3 
               
               
                 ″ 
                 −10 
                 240 
                 14.2 
                 99.6 
                 74.0 
                 10.5 
               
               
                 ″ 
                 0 
                 60 
                 12.5 
                 99.6 
                 74.0 
                 9.3 
               
               
                   
               
               
                 Key: heptane = n-heptane; 3MHx = 3-methyl hexane; 23DMP = 2,3-dimethyl pentane; 24DMP = 2,4-dimethyl pentane; conv. = conversion; res. = residence  
               
               
                 As can be seen from the data, the selectivity of the process towards triptane is at least 8.0% under the reaction conditions are employed.
A process for the production of triptane, said process comprising: isomerising a hydrocarbon feedstock by containing said feedstock with an isomerisation catalyst at a reaction temperature of -50 to 25deg C., and a contact time of 0.01 to 150 hours, such that the triptane selectivety of the isomerisation reaction is at least 5% as a proportion of said hydrocarbon feedstock.