EP0781831A1 - Process for lowering the content of benzene and of light unsaturated compounds in hydrocarbon fractions - Google Patents

Abstract

The treatment of a charge comprising mainly 5C hydrocarbons and containing at least one unsaturated 6C maximum compound including benzene comprises treating the charge in a distillation column consisting of a draining zone and a rectification zone associated with a hydrogenation reaction zone, comprising at least one catalytic bed, in which part of the unsaturated compounds are hydrogenated in the presence of a catalyst and a hydrogen (H2) flow; the charge from the reaction zone is sampled at a suitable level and returned to the distillation zone to ensure continuity of the process. From the head and base of the column are obtained effluents depleted in unsaturated compounds up to 6C; part of the head effluent (comprising 5 and/or 6C paraffins) is treated in an isomerisation zone in the presence of a catalyst to give an isomerate.

Description

The invention relates to a process for the selective reduction of the content of light unsaturated compounds (that is to say containing at most six carbon atoms per molecule) including benzene, of a hydrocarbon fraction essentially comprising at least 5 carbon atoms per molecule, without appreciable loss of the octane number, said process comprising passing said cut through a distillation zone associated with a hydrogenation reaction zone, followed by the passage of part of the effluent of the distillation zone mainly comprising C ₅ -C ₆ hydrocarbons, that is to say containing 5 and / or 6 carbon atoms per molecule, in a zone for isomerization of paraffins.

Given the recognized harmfulness of benzene and olefins, unsaturated compounds, the general tendency is to reduce the content of these constituents in gasolines.

Benzene has carcinogenic properties and it is therefore required to limit as much as possible any possibility of polluting the ambient air, in particular by practically excluding it from automotive fuels. In the United States reformulated fuels must not contain more than 1% benzene; in Europe, even if the specifications are not yet as strict, it is recommended to gradually move towards this value.

Olefins have been recognized as being among the most reactive hydrocarbons in the photochemical reaction cycle with nitrogen oxides, which occurs in the atmosphere and which leads to the formation of ozone. An increase in the concentration of ozone in the air can cause respiratory problems. The reduction in the olefin content of gasolines, and more particularly of the lighter olefins which have the most tendency to volatilize during handling of the fuel, is therefore desirable.

The benzene content of a gasoline is very largely dependent on that of the reformate component of this gasoline. The reformate results from a catalytic treatment of naphtha intended to produce aromatic hydrocarbons, mainly comprising from 6 to 9 carbon atoms in their molecule and whose very high octane number gives the gasoline its anti-knock properties.

For the harmfulness reasons described above, it is therefore necessary to reduce the benzene content of the reformate as much as possible. Several routes are possible.

A first route consists in limiting, in the naphtha constituting the charge of a catalytic reforming unit, the content of benzene precursors, such as cyclohexane and methylcyclopentane. This solution effectively makes it possible to significantly reduce the benzene content of the effluent from the reforming unit but cannot be sufficient in itself when it comes to descending to contents as low as 1%. A second way consists in eliminating, by distillation, a light fraction of the reformate containing benzene. This solution leads to a loss of around 15 to 20% of hydrocarbons which would be recoverable in gasoline. A third way consists in extracting the benzene present in the effluent from the reforming unit. Several known techniques are in principle applicable: solvent extraction, extractive distillation, adsorption. None of these techniques are applied industrially, because none of them allows to selectively extract benzene in an economical manner. A fourth way consists in chemically transforming benzene to convert it into a constituent not covered by legal limitations. Alkylation with ethylene, for example, transforms benzene mainly into ethylbenzene. This operation is however expensive because of the intervention of secondary reactions which require costly separations in energy.

Benzene from a reformate can also be hydrogenated to cyclohexane. As it is impossible to selectively hydrogenate benzene from a mixture of hydrocarbons also containing toluene and xylenes, it is therefore necessary to fractionate this mixture beforehand so as to isolate a cut containing only benzene, which can then be hydrogen. A process has also been described in which the benzene hydrogenation catalyst is included in the rectification zone of the distillation column which separates benzene from the other aromatics (Benzene Reduction - Kerry Rock and Gary Gildert CDTECH - 1994 Conference on Clean Air Act Implementation and Reformulated Gasoline - Oct. 94), which saves on equipment.

The hydrogenation of benzene from a reformate leads to a loss of octane number. This loss of octane can be compensated for by the addition of compounds with a high octane number, for example ethers such as MTBE or ETBE, or branched paraffinic hydrocarbons. These branched paraffinic hydrocarbons can be generated from the reformate itself, by isomerization of the linear paraffins. However, it is known that the catalysts for isomerization of linear paraffins into branched paraffins are not inactive with respect to hydrocarbons of other chemical families. Among those which distill with benzene due to the azeotropy phenomenon, cyclohexane for example is partially converted into methylcyclopentane. This reaction of naphthenic products competes on the catalyst with the reaction of isomerization of paraffins and consequently reduces their progress. On the other hand, isoparaffins with 7 carbon atoms per molecule undergo cracking which leads on the one hand to a progressive fouling of the isomerization catalyst, therefore to a lower activity, and on the other hand, to a reduction in the yield of sought-after product, that is to say a light reformate to be included in petrol.

The process according to the invention avoids the disadvantages mentioned, that is to say that it makes it possible to produce, at the lowest cost, from a raw reformate, a reformate depleted in benzene or, if necessary, almost completely purified from benzene as well as other unsaturated hydrocarbons containing at most six carbon atoms per molecule such as light olefins, without significant loss of yield, and with very little loss or with an increase in octane number.

The process is characterized by the integration of the three distillation, hydrogenation and isomerization operations arranged and operated so as to avoid at least in part, preferably in the main, entrainment, by azeotropy with benzene, of cyclohexane and isoparaffins with 7 carbon atoms per molecule, in the distillate which is directed to isomerization. Thus, the process according to the invention at least partially performs the selective hydrogenation of benzene and of any unsaturated compound comprising at most six carbon atoms per molecule and different from benzene, possibly present in the feed.

• on traite ladite charge dans une zone de distillation, comportant une zone d'épuisement et une zone de rectification, associée à une zone réactionnelle d'hydrogénation, comportant au moins un lit catalytique, dans laquelle on réalise l'hydrogénation d'au moins une partie des composés insaturés comprenant au plus six atomes de carbone par molécule, c'est-à-dire comprenant jusqu'à six (inclus) atomes de carbone par molécule, et contenus dans la charge, en présence d'un catalyseur d'hydrogénation et d'un flux gazeux comprenant, de préférence en majeure partie, de l'hydrogène, la charge de la zone réactionnelle étant prélevée à la hauteur d'un niveau de prélèvement et représentant au moins une partie, de préférence la majeure partie, du liquide coulant dans la zone de distillation, de préférence coulant dans la zone de rectification et de façon encore plus préférée coulant à un niveau intermédiaire de la zone de rectification, l'effluent de la zone réactionnelle étant au moins en partie, de préférence en majeure partie, réintroduit dans la zone de distillation, de manière à assurer la continuité de la distillation, et de façon à sortir finalement en tête de la zone de distillation un effluent très appauvri en composés insaturés comprenant au plus six atomes de carbone par molécule, et en fond de zone de distillation un effluent également appauvri en composés insaturés comprenant au plus six atomes de carbone par molécule.
• on traite dans une zone d'isomérisation au moins une partie, de préférence la majeure partie, de l'effluent soutiré en tête de zone de distillation, ladite partie renfermant des paraffines contenant 5 et/ou 6 atomes de carbone par molécule (c'est-à-dire choisies dans le groupe formé par les paraffines comportant 5 atomes de carbone par molécule et les paraffines comportant 6 atomes de carbone par molécule), éventuellement en présence d'une autre coupe comprenant des paraffines renfermant en majeure partie 5 et/ou 6 atomes de carbones par molécule, en présence d'un catalyseur d'isomérisation, de façon à obtenir un isomérat.

The process according to the invention is a process for treating a feedstock, consisting mainly of hydrocarbons comprising at least 5, preferably between 5 and 9 carbon atoms per molecule, and comprising at least one unsaturated compound containing at most six carbon atoms per molecule including benzene, such as:

• said feed is treated in a distillation zone, comprising a depletion zone and a rectification zone, associated with a hydrogenation reaction zone, comprising at least one catalytic bed, in which the hydrogenation of at least one is carried out part of the unsaturated compounds comprising at most six carbon atoms per molecule, that is to say comprising up to six (inclusive) carbon atoms per molecule, and contained in the feed, in the presence of a hydrogenation catalyst and of a gas flow preferably comprising mainly hydrogen, the charge of the reaction zone being taken off at the level of a taking off level and representing at least a part, preferably the major part, of the liquid flowing in the distillation zone, preferably flowing in the rectification zone and even more preferably flowing at an intermediate level of the rectification zone, the effluent from the reaction zone tional being at least partly, preferably mainly, reintroduced into the distillation zone, so as to ensure the continuity of the distillation, and so as to finally leave at the top of the distillation zone an effluent very depleted in unsaturated compounds comprising at most six carbon atoms per molecule, and at the bottom of the distillation zone an effluent also depleted in unsaturated compounds comprising at most six carbon atoms per molecule.
• at least part, preferably most, of the effluent withdrawn at the head of the distillation zone is treated in an isomerization zone, said part containing paraffins containing 5 and / or 6 carbon atoms per molecule (c ' that is to say chosen from the group formed by paraffins containing 5 carbon atoms per molecule and paraffins comprising 6 carbon atoms per molecule), optionally in the presence of another cut comprising paraffins containing mainly 5 and / or 6 carbon atoms per molecule, in the presence of an isomerization catalyst, so as to obtain an isomerate.

The other cut comprising paraffins containing mainly 5 and / or 6 carbon atoms per molecule, possibly present in the isomerization charge with the part of the effluent withdrawn at the head of the distillation zone, comes from any source known to those skilled in the art. We can cite as an indication a so-called light naphtha cut coming from a naphtha fractionation unit.

The charge which feeds the distillation zone is introduced into said zone generally at least at one level of said zone, preferably mainly at one level of said zone.

The distillation zone generally comprises at least one column provided with at least one internal distillation chosen from the group formed by trays, bulk packings and structured packings, as is known to those skilled in the art, such that the total overall efficiency is generally at least equal to five theoretical stages. In the cases known to those skilled in the art where the implementation of a single column poses problems, it is generally preferred to divide said zone so as to ultimately use at least two columns which, placed end to end, form said zone, that is to say that the zones of rectification, possibly reaction and exhaustion are distributed over the columns. In practice, when the reaction zone is at least partially internal to the distillation zone, the rectification zone or the exhaustion zone, and preferably the exhaustion zone, can generally be located in at least one column different from the column comprising the internal part of the reaction zone.

The hydrogenation reaction zone generally comprises at least one catalytic hydrogenation bed, preferably from 1 to 4 catalytic bed (s); in the case where at least two catalytic beds are incorporated in the distillation zone, these two beds are optionally separated by at least one internal distillation. The hydrogenation reaction zone at least partially performs the hydrogenation of the benzene present in the feed, generally in such a way that the benzene content of the overhead effluent is at most equal to a certain content, and said reaction zone performs at less in part, preferably in major part, the hydrogenation of any unsaturated compound comprising at most six carbon atoms per molecule and different from benzene, possibly present in the feed.

According to a first embodiment of the invention, the process according to the invention is such that the hydrogenation reaction zone is at least in part, of preferably entirely, internal to the distillation zone. Then, for the part of the reaction zone internal to the distillation zone, the withdrawal of liquid is done naturally by flow in the part of the reaction zone internal to the distillation zone, and the reintroduction of the effluent in the distillation zone is also done naturally by flow of the liquid from the reaction zone internal to the distillation zone so as to ensure the continuity of the distillation. In addition, the process according to the invention is preferably such that the flow of the liquid to be hydrogenated is co-current with the flow of the gas flow comprising hydrogen, for any catalytic bed of the internal part of the zone d hydrogenation, and even more preferably such that the flow of the liquid to be hydrogenated is co-current with the flow of the gas stream comprising hydrogen and such that the distillation vapor is separated from said liquid, for any catalytic bed of the internal part of the hydrogenation zone.

According to a second embodiment of the invention, independently of the previous embodiment, the process according to the invention is such that the hydrogenation reaction zone is at least partly, preferably entirely, external to the distillation zone . Then the effluent from at least one catalytic bed from the external part of the dehydrogenation zone is generally reintroduced substantially near a level of sampling, preferably the level of sampling which fed said catalytic bed. Generally, the method according to the invention comprises from 1 to 4 sampling level (s) which feed (s) the external part of the hydrogenation zone. So two cases can arise. In the first case, the external part of the hydrogenation zone is supplied by a single level of sampling, and then, if said part comprises at least two catalytic beds distributed in at least two reactors, said reactors are arranged in series or in parallel. In the second case, preferred according to the present invention, the external part of the hydrogenation zone is supplied by at least two sampling levels.

According to a third embodiment of the invention, which combines the two embodiments described above, the method according to the invention is such that the hydrogenation zone is both partially incorporated in the distillation zone, it is that is to say internal to the distillation zone, and partially external to the distillation zone. According to such an embodiment, the hydrogenation zone comprises at least two catalytic beds, at least one catalytic bed being internal to the distillation zone, and at least one other catalytic bed being external to the distillation zone. In the case where the external part of the hydrogenation zone comprises at least two catalytic beds, each catalytic bed is supplied by a single level of sampling, preferably associated with a single level where the effluent of said catalytic bed from the external part of the hydrogenation zone is reintroduced, said level of sampling being distinct from the level of sampling which feeds the other catalytic bed (s). Generally, the liquid to be hydrogenated, either partially or completely, first circulates in the external part of the hydrogenation zone and then the internal part of said zone. The part of the reaction zone internal to the distillation zone has the characteristics described in the first embodiment. The part of the reaction zone external to the distillation zone has the characteristics described in the second embodiment.

According to another embodiment of the invention, independently or not of the preceding embodiments, the method according to the invention is such that the flow of the liquid to be hydrogenated is co-current or counter-current, preferably co-current , to the flow of the gas stream comprising hydrogen, for any catalytic bed in the hydrogenation zone.

For carrying out the hydrogenation according to the process of the invention, the theoretical molar ratio of hydrogen necessary for the desired conversion of benzene is 3. The quantity of hydrogen distributed, in the gas flow, before or in the zone of hydrogenation is optionally in excess with respect to this stoichiometry, and all the more so since it is necessary to hydrogenate, in addition to the benzene present in the feed, at least partially any unsaturated compound comprising at most six carbon atoms per molecule and present in said charge. The excess hydrogen, if it exists, can be advantageously recovered, for example according to one of the techniques described below. According to a first technique, the excess hydrogen which leaves the top of the distillation zone is recovered, then compressed and reused in the hydrogenation zone. According to a second technique, the excess hydrogen which leaves the top of the distillation zone is recovered, then compressed and reused in the isomerization zone. According to a third technique, the excess hydrogen which leaves the top of the distillation zone is recovered, then injected upstream of the compression stages associated with a catalytic reforming unit, in mixture with hydrogen originating from said unit, said unit preferably operating at low pressure, that is to say generally a pressure of less than 8 bar (1 bar = 10 ⁵ Pa).

The hydrogen used according to the invention for the hydrogenation of unsaturated compounds comprising at most six carbon atoms per molecule, and included in the gas stream, can come from all sources producing hydrogen at least 50% volume of purity , preferably at least 80% volume of purity and even more preferably at least 90% volume of purity. For example, mention may be made of hydrogen from catalytic reforming processes, methanation, P.S.A. (alternating pressure adsorption), electrochemical generation, steam cracking or steam reforming. It is also conceivable, for example, that the hydrogen injected into the hydrogenation process first passes through the isomerization step. In such a case, hydrogen is injected into the isomerization unit to delay the deactivation of the isomerization catalyst by carbon deposition. Unconsumed hydrogen from the isomerization zone can then be purified and then used in the hydrogenation unit.

One of the preferred embodiments of the process according to the invention, independent or not of the preceding embodiments, is such that the bottom effluent from the distillation zone is mixed at least in part with the isomerization effluent. The mixture thus obtained can, after optional stabilization, be used as fuel either directly or by incorporation into the fuel fractions.

Generally, preferably, the operating conditions are judiciously chosen, in relation to the nature of the feed and with other parameters known to the specialist in reactive distillation, such as the distillate / feed ratio, so that the effluent at the top of the distillation zone is practically free of cyclohexane and isoparaffins comprising 7 carbon atoms per molecule. Thus, the process according to the invention is generally and preferably such that the overhead effluent from the distillation zone is practically free from cyclohexane and isoparaffins comprising 7 carbon atoms per molecule.

When the hydrogenation zone is at least partly incorporated into the distillation zone, the hydrogenation catalyst can be placed in said incorporated part according to the various technologies proposed for conducting catalytic distillations. They are essentially of two types.

According to the first type of technology, the reaction and the distillation proceed simultaneously in the same physical space, as teach for example the patent application WO-A-90 / 02.603, the patents US-A-4,471,154, US- A-4.475.005, US-A-4.215.011, US-A-4.307.254, US-A-4.336.407, US-A-4.439.350, US-A-5.189.001, US-A- 5.266.546, US-A-5.073.236, US-A-5.215.011, US-A-5.275.790, US-A-5.338.517, US-A-5.308.592, US-A-5.236. 663, US-A-5,338,518, as well as the patents EP-B1-0.008.860, EP-B1-0.448.884, EP-B1-0.396.650 and EP-B1-0.494.550 and the patent application EP-A1-0.559.511. The catalyst is then generally in contact with a descending liquid phase, generated by the reflux introduced at the top of the distillation zone, and with an ascending vapor phase, generated by the reboiling vapor introduced at the bottom of the zone. According to this type of technology, the gas stream comprising hydrogen necessary for the reaction zone, for carrying out the process according to the invention, could be joined to the vapor phase, substantially at the inlet of at least one catalytic bed of the reaction zone.

According to the second type of technology, the catalyst is arranged in such a way that the reaction and the distillation generally proceed independently and consecutively, as taught for example by the patents US-A-4,847,430, US-A-5,130. 102 and US-A-5,368,691, the vapor from the distillation zone practically not passing through any catalytic bed of the reaction zone. Thus the process according to the invention is generally such that the flow of the liquid to be hydrogenated is co-current with the flow of the gas flow comprising hydrogen and such that the distillation vapor is practically not in contact with the catalyst (which generally results in practice in that said vapor is separated from said liquid to be hydrogenated), for any catalytic bed of the internal part of the hydrogenation zone. In all cases of this second type of technology, any catalytic bed of the part of the reaction zone which is in the distillation zone is generally such that the gas flow comprising hydrogen and the flow of the liquid which will react circulate at co-current, generally ascending, through said bed, even if overall, in the catalytic distillation zone, the gas flow comprising hydrogen and the flow of the liquid which will react circulate against the current. Such systems generally include at least one liquid distribution device which can, for example, be a liquid distributor, in any catalytic bed in the reaction zone. However, since these technologies were designed for reactions catalytic intervening between liquid reactants, they cannot be suitable without modification for a catalytic hydrogenation reaction, for which one of the reactants, hydrogen, is in the gaseous state. For any catalytic bed in the internal part of the hydrogenation zone, it is therefore generally necessary to add a gas flow distribution device comprising hydrogen, for example according to one of the three techniques described below. Thus, the internal part of the hydrogenation zone comprises at least one liquid distribution device and at least one gas flow distribution device comprising hydrogen, for any catalytic bed of the internal part of the hydrogenation zone . According to a first technique, the device for distributing the gas flow comprising hydrogen is placed before the liquid dispensing device, and therefore before the catalytic bed. According to a second technique, the device for distributing the gas flow comprising hydrogen is arranged at the level of the device for distributing liquid, in such a way that the gas flow comprising hydrogen is introduced into the liquid before the catalytic bed. According to a third technique, the gas flow distribution device comprising hydrogen is placed after the liquid distribution device, and therefore within the catalytic bed, preferably not far from said liquid distribution device in said catalytic bed. The terms "before" and "after" used above are understood with respect to the direction of circulation of the liquid which will pass through the catalytic bed, that is to say generally in the upward direction.

One of the preferred embodiments of the process according to the invention is such that the catalyst of the internal part of the hydrogenation zone is placed in the reaction zone according to the basic device described in patent US-A-5,368,691, fitted with so that any catalytic bed internal to the distillation zone is fed by a gas flow comprising hydrogen, regularly distributed at its base, for example according to one of the three techniques described above. According to this technology, if the distillation zone comprises a single column and if the hydrogenation zone is entirely internal to said column, the catalyst included in any catalytic bed, internal to the distillation zone, is then in contact with a phase. ascending liquid, generated by the reflux introduced at the top of the distillation column, and with the gas flow comprising hydrogen which circulates in the same direction as the liquid; contact with the vapor phase of the distillation is avoided by passing it through at least one specially fitted chimney.

When the hydrogenation zone is at least partly internal to the distillation zone, the operating conditions of the part of the hydrogenation zone internal to the distillation zone are linked to the operating conditions of the distillation. The distillation is carried out so that its bottom product contains the major part of the cyclohexane and isoparaffins with 7 carbon atoms of the feed, as well as of the cyclohexane formed by hydrogenation of benzene. It is carried out under a pressure generally between 2 and 20 bar, preferably between 4 and 10 bar (1 bar = 10 ⁵ Pa), with a reflux rate between 1 and 10, and preferably between 3 and 6. The zone head temperature is generally between 40 and 180 ° C. and the zone bottom temperature is generally between 120 and 280 ° C. The hydrogenation reaction is carried out under conditions which are most generally intermediate between those established at the head and at the bottom of the distillation zone, at a temperature between 100 and 200 ° C., and preferably between 120 and 180 ° C. , and at a pressure between 2 and 20 bar, preferably between 4 and 10 bar. The liquid subjected to hydrogenation is supplied by a gas stream comprising hydrogen, the flow rate of which depends on the concentration of benzene in said liquid and, more generally, unsaturated compounds containing at most six carbon atoms per molecule of the charge. from the distillation zone. It is generally at least equal to the flow rate corresponding to the stoichiometry of the hydrogenation reactions involved (hydrogenation of benzene and other unsaturated compounds containing at most six carbon atoms per molecule, included in the hydrogenation charge) and at most equal at the flow rate corresponding to 10 times the stoichiometry, preferably between 1 and 6 times the stoichiometry, even more preferably between 1 and 3 times the stoichiometry.

When the hydrogenation zone is partly external to the distillation zone, the catalyst placed in said external part is according to any technology known to those skilled in the art under independent operating conditions (temperature, pressure, etc.) or no, preferably independent, of the operating conditions of the distillation zone.

In the part of the hydrogenation zone external to the distillation zone, the operating conditions are generally as follows. The pressure required for this hydrogenation step is generally between 1 and 60 bar absolute, preferably between 2 and 50 bar and even more preferably between 5 and 35 bar.

The operating temperature of the external part of the hydrogenation zone is generally between 100 and 400 ° C, preferably between 120 and 350 ° C and preferably between 140 and 320 ° C. The space velocity within the external part of said hydrogenation zone, calculated with respect to the catalyst, is generally between 1 and 50 and more particularly between 1 and 30 h ^-1 (volume of charge per volume of catalyst and per hour ). The hydrogen flow rate corresponding to the stoichiometry of the hydrogenation reactions involved is between 0.5 and 10 times said stoichiometry, preferably between 1 and 6 times said stoichiometry and even more preferably between 1 and 3 times said stoichiometry . However, the temperature and pressure conditions can also, within the scope of the process of the present invention, be between those which are established at the top and at the bottom of the distillation zone.

More generally, whatever the position of the hydrogenation zone relative to the distillation zone, the catalyst used in the hydrogenation zone according to the process of the present invention generally comprises at least one metal chosen from the group formed by nickel and platinum, used as it is or preferably deposited on a support. The metal must generally be in reduced form at least for 50% by weight of its whole. However, any other hydrogenation catalyst known to those skilled in the art can also be chosen.

When using platinum, the catalyst can advantageously contain at least one halogen in a proportion by weight relative to the catalyst of between 0.2 and 2%. Preferably, chlorine or fluorine or a combination of the two is used in a proportion relative to the total weight of catalyst of between 0.2 and 1.5%. In the case of the use of a platinum-containing catalyst, a catalyst is generally used such that the average size of the platinum crystallites is less than 60.10 ^-10 m, preferably less than 20.10 ^-10 m, even more so preferred less than 10.10 ^-10 m. In addition, the total proportion of platinum relative to the total weight of catalyst is generally between 0.1 and 1% and preferably between 0.1 and 0.6%.

In the case of the use of nickel, the proportion of nickel relative to the total weight of catalyst is between 5 and 70%, more particularly between 10 and 70% and preferably between 15 and 65%. In addition, one generally uses a catalyst such that the average size of the nickel crystallites is less than 100.10 ^-10 m, preferably less than 80.10 ^-10 m, even more preferably less than 60.10 ¹⁰ m.

The support is generally chosen from the group formed by alumina, silica-aluminas, silica, zeolites, activated carbon, clays, aluminous cements, rare earth oxides and alkaline earth oxides, alone or in mixture. A support based on alumina or silica is preferably used, with a specific surface area of between 30 and 300 m ² / g, preferably between 90 and 260 m ² / g.

The isomerization catalyst used in the isomerization zone according to the present invention is generally of two types. However, any other isomerization catalyst known to a person skilled in the art can also be chosen.

The first type of catalyst is based on alumina. Preferably, it comprises at least one metal from group VIII of the periodic table of the elements and a support comprising alumina. Preferably, it also comprises at least one halogen, preferably chlorine. Thus a preferred catalyst according to the present invention comprises at least one group VIII metal deposited on a support consisting of eta alumina and / or gamma alumina, that is to say that for example said support consists of alumina eta and gamma alumina, the alumina eta content being between 85 and 95% by weight relative to the support, preferably between 88 and 92% by weight, and even more preferably between 89 and 91% by weight, the complement to 100% weight of the support consisting of gamma alumina. However, the catalyst support can also, for example, consist essentially of gamma alumina. The group VIII metal is preferably chosen from the group formed by platinum, palladium and nickel.

The alumina was optionally used in the present invention has a specific surface generally between 400 and 600 m ² / g and preferably between 420 and 550 m ² / g, and a total pore volume generally between 0.3 and 0 , 5 cm ³ / g and preferably between 0.35 and 0.45 cm ³ / g.

The gamma alumina optionally used in the present invention generally has a specific surface of between 150 and 300 m ² / g and of preferably between 180 and 250 m ² / g, a total pore volume generally between 0.4 and 0.8 cm ³ / g and preferably between 0.45 and 0.7 cm ³ / g.

The two types of alumina, when used as a mixture, are mixed and shaped, in proportions defined by any technique known to those skilled in the art, for example by extrusion through a die, by pelleting. or coating.

A second type of catalyst used in the isomerization zone according to the process of the present invention is a zeolite-based catalyst, that is to say comprising at least one group VIII metal and a zeolite. Different zeolites can be used for said catalyst; said zeolite is preferably chosen from the group formed by mordenite or omega Ω zeolite. A mordenite having a Si / Al (atomic) ratio between 5 and 50 and preferably between 5 and 30 is preferably used. sodium less than 0.2% and preferably less than 0.1% (relative to the weight of dry zeolite), a volume of mesh V of the elementary mesh between 2.78 and 2.73 nm ³ and so preferred between 2.77 and 2.74 nm ³ , a benzene absorption capacity greater than 5% and preferably greater than 8% (relative to the weight of dry solid). The mordenite thus prepared is then mixed with a generally amorphous matrix (alumina, silica alumina, kaolin, ...) and shaped by any method known to those skilled in the art (extrusion, pelletizing, coating). The mordenite content of the support thus obtained must be greater than 40% and preferably greater than 60% by weight.

It is also possible to use a catalyst based on an omega Ω or mazzite zeolite. Said zeolite has a SiO ₂ / Al ₂ O ₃ molar ratio of between 6.5 and 80, preferably between 10 and 40, a sodium content by weight of less than 0.2%, preferably of less than 0.1%, per relative to the weight of dry zeolite. It usually has crystalline parameters "a" and "c" respectively less than or equal to 1.814 nm and 0.760 nm (1 nm = 10 ^-9 m), preferably respectively between 1.814 and 1.794 nm and between 0.760 and 0.749 nm, a nitrogen adsorption capacity, measured at 77 K under a partial pressure equal to 0.19 bar, greater than approximately 8% by weight, preferably greater than approximately 11% by weight. Its porous distribution generally comprises between 5 and 50% of the porous volume contained in radius pores (measured by the BJH method) located between 1.5 and 14 nm, preferably between 2.0 and 8.0 nm (mesopores). Generally speaking, its DX crystallinity level (measured according to its X-ray diffractogram) is greater than 60%.

The zeolitic support thus obtained has a specific surface generally between 300 and 550 m ² / g and preferably between 350 and 500 m ² / g and a pore volume generally between 0.3 and 0.6 cm ³ / g and of preferably between 0.35 and 0.5 cm ³ / g.

Whatever the support of the isomerization catalyst (alumina or zeolite), at least one hydrogenating metal from group VIII, preferably chosen from the group formed by platinum, palladium and nickel, is then deposited on this support, by any technique known to those skilled in the art, for example in the case of platinum by anion exchange in the form of hexachloroplatinic acid when the support is alumina and by cation exchange with platinum tetramine chloride when the support is a zeolite.

In the case of platinum or palladium, the content by weight is between 0.05 and 1% and preferably between 0.1 and 0.6%. In the case of nickel, the weight content is between 0.1 and 10% and preferably between 0.2 and 5%.

The isomerization catalyst thus prepared can be reduced under hydrogen. In the case where the support is based on alumina, said catalyst is subjected to a halogenation treatment, preferably of chlorination, by any halogenated compound, preferably chlorinated, known to a person skilled in the art such as for example carbon tetrachloride or perchlorethylene. The halogen, preferably chlorine, content of the final catalyst is preferably between 5 and 15% by weight and preferably between 6 and 12% by weight. This halogenation treatment, preferably chlorination, of the catalyst can be carried out either directly in the unit before injection of the charge ("in-situ") or off site. In such a case, it is also possible to carry out the halogenation treatment, preferably chlorination, prior to the reduction treatment of the catalyst under hydrogen.

The operating conditions used in the isomerization zone are generally those described below, depending on the type of catalyst.

With the first type of catalyst, based on alumina, the temperature is generally between 80 and 300 ° C. and preferably between 100 and 200 ° C. The partial pressure of hydrogen is between 0.1 and 70 bar and preferably between 1 and 50 bar. The space velocity is between 0.2 and 10, preferably between 0.5 and 5, liters of liquid hydrocarbons per liter of catalyst and per hour. The hydrogen to hydrocarbons molar ratio at the inlet of the isomerization zone is such that the hydrogen to hydrocarbons molar ratio in the isomerate is greater than 0.06 and preferably between 0.06 and 10.

With the second type of catalyst, zeolitic, the temperature is generally between 200 and 300 ° C and preferably between 230 and 280 ° C, and the partial pressure of hydrogen is between 0.1 and 70 bar and preferably between 1 and 50 bar. The space velocity is generally between 0.5 and 10, preferably between 1 and 5 liters of liquid hydrocarbons per liter of catalyst and per hour. The hydrogen to hydrocarbons molar ratio in the isomerate can vary between wide limits and is generally between 0.07 and 15 and preferably between 1 and 5.

Figures 1 to 3 each constitute an illustration of a possibility of carrying out the method according to the invention. Similar devices are represented by the same numbers in all the figures.

A first embodiment of the process is shown in FIG. 1. The crude C ₅ ⁺ reformate, generally containing small quantities of C ₄ ^- hydrocarbons, is sent to a column 2 by line 1. Said column contains distillation internals, which are for example in the case shown in Figure 1 of the plates or the lining, represented in part by dotted lines in said figure. It also contains at least one internal catalytic 3 containing a hydrogenation catalyst, which can be alternated with the internal distillation. The catalytic internals are supplied at their base, by lines 4c and 4d, by hydrogen coming from lines 4, then 4a and 4b. At the bottom of the column, the least volatile fraction of the reformate, consisting mainly of hydrocarbons with 7 carbon atoms and more, is recovered by line 5, reboiled in exchanger 6 and evacuated by line 7. Steam reboiling is reintroduced into the column by line 8. At the top of the column, the vapor of light hydrocarbons, that is to say comprising mainly 6 carbon atoms and less per molecule, is sent by line 9 into a condenser 10 then in a balloon 11 where there is a separation between a liquid phase and a vapor phase mainly consisting of excess hydrogen possibly sent by lines 16 then 4a then 4b then 4c or 4d.

The vapor phase is evacuated from the flask by lines 14 and then 15. A fraction is optionally recycled to the column by line 16, after being re-pressurized by means of a device not shown in FIG. 1.

The liquid phase of the flask 11 is partly returned, by line 12, to the column head to ensure reflux. The other part is directed by lines 13 then 17 to the isomerization reactor 18. A stream of hydrogen is optionally added thereto by lines 4 then 4a. The isomerate is recovered by line 19, cooled, and sent to a flask 20 where a vapor phase consisting essentially of hydrogen separates, which is evacuated by lines 22 then 23, and optionally recycled after purification to the hydrogen circuit by line 24 then by lines 4a, 4b, and 4c or 4d.

The liquid phase is drawn off through line 21 and constitutes, after stabilization if necessary, a component for gasolines, almost free of unsaturated compounds comprising at most 6 carbon atoms per molecule, of high octane number.

According to a second embodiment of the process, represented in FIG. 2, the crude C ₅ ⁺ reformate, generally containing small quantities of C ₄ ^- hydrocarbons, is sent by line 1 to a distillation column 2, provided with distillation internals which are, for example in the case of FIG. 2, distillation plates, as well as a withdrawal (or sampling) plate of liquid phase. The liquid phase withdrawn from the withdrawal plate by line 25 is brought into contact with hydrogen supplied by lines 4, 4a and 4b, and directed to a hydrogenation reactor 33. The hydrogenation reactor can operate either by upward flow, ie downward flow as shown in FIG. 2. The effluent from this reactor is recovered by line 26 and recycled to the distillation column by lines 27 then 32, generally in the upper part of the distillation zone located under the racking plate near said plate. It is generally considered that a maximum of four hydrogenation reactors can constitute the hydrogenation zone, in the case where it is external to the distillation zone, regardless of the number of levels of sampling (x).

According to a variant of the process, all or part of the reactor effluent recovered by line 26 is cooled (exchanger not shown) and directed by line 28 to the tank 29 where a vapor phase rich in hydrogen separates, evacuated by the line 30, and a liquid phase which is recycled to column 2 by lines 31 and 32. The effluents from the top and bottom of the column are treated as described above for the first embodiment of the process.

According to a third version of the process, represented in FIG. 3, the hydrogenation zone is shared between a part internal to the distillation column, as described for the first version of the process, and a part external to this column, as described for the second version of the process.

EXAMPLES

The examples which follow illustrate the invention in the particular case of FIG. 1.

Example 1 :

A metal distillation column with a diameter of 50 mm is used, made adiabatic by heating jackets whose temperatures are regulated so as to reproduce the temperature gradient which is established in the column. At a height of 4.5 m, the column includes, from head to toe: a rectification zone composed of 11 perforated trays with overflow and descent, a hydrogenating catalytic distillation zone and an exhaustion zone composed of 63 trays perforated. The hydrogenating catalytic distillation zone consists of three catalytic distillation doublets, each doublet itself being constituted by a catalytic cell surmounted by three perforated plates. The construction detail of a catalytic cell as well as its arrangement in the column are shown schematically for information in Figure 4. The catalytic cell 41 consists of a cylindrical container with a flat bottom, with an outside diameter less than 2 mm in lower diameter of the column. It is provided at its lower part, above the bottom, with a grid 42 which serves both as a support for the catalyst and as a liquid distributor for hydrogen, and at its upper part, with a retaining grid catalyst 43, the height of which can be varied. The catalyst 44 fills the entire volume between these two grids. The catalytic cell receives the liquid coming from the upper distillation plate 45, by the descent 46. After having traversed the cell in the upward direction, the liquid is evacuated by overflow by the descent 47 and flows on the lower distillation plate 48. The steam from the lower plate 48 borrows the central chimney 49 secured to the cell, entering through orifices 50 (only one visible in the figure) and emerging under the upper plate 45 through orifices 51 (only one visible in the figure) . Hydrogen is introduced at the foot of the catalytic cell through the tube 52, then through the orifices 53 (six in total) distributed over the periphery of the cell, in the immediate vicinity of the bottom. Seals 54 prevent any hydrogen leakage before its arrival on the catalytic bed.

Each of the three cells is packed with 36 g of nickel catalyst sold by the company PROCATALYSE under the reference LD 746. On the 37th tray of the column, starting from the bottom, 250 g / h of a reformate consisting essentially of hydrocarbons having at least 5 carbon atoms in their molecule, the composition of which is presented in the second column of Table 1. A flow rate of 4.5 Nl / h of hydrogen is also introduced at the base of each cell. The column is brought into operation by establishing a reflux rate equal to 5 and by regulating the bottom temperature at 195 ° C. and the absolute pressure at 6 bar.

In a steady state, a residue and a distillate are collected at the rate of 181 g / h and 69 g / h, respectively, the compositions of which are given in the third and fourth columns of Table 1.

The distillate is sent together with hydrogen, with a hydrogen / hydrocarbon molar ratio fixed at 0.125, in an isomerization reactor containing 57 g of catalyst based on platinum on chlorinated alumina, sold by the company PROCATALYSE under the reference IS612A , operating at a temperature of 150 ° C and a pressure of 30 bar. The effluent from the isomerization reactor or isomerate has the composition presented in the last column of Table 1.

The last three lines of table 1 show the octane numbers RON (Research), MON (Engine) and (RON + MON) / 2 (Medium octane) of the reformate, column effluents and isomerate . The isomerate has an octane number by 3 points higher than the distillate, and can be valued as a fuel component, provided that it is stabilized, that is to say, that it is rid by distillation of the 3% of very volatile constituents (C ₃ ^- ) formed during isomerization, mainly by decomposition of isoparaffins with 7 carbon atoms per molecule. By mixing the residue from the distillation with the stabilized isomerate, a gasoline almost free of benzene and olefins with an average octane number equal to 90.3 is reconstituted. Compared with the starting reformate, the reconstituted petrol therefore shows an average octane loss of 0.3 points and is produced with a yield loss of 0.8 points. Table 1: compositions (% by weight) and octane numbers of the different streams for example 1 Reformate Residue Distillate Isomerate Hydrocarbons C6 ^- 26.4 0.20 94.9 97.9 of which: C3- - - - 3.0 olefins 0.19 - - - benzene 4.70 - 0.48 - cyclohexane 0.08 0.19 16.3 6.85 Hydrocarbons C7 + 73.6 99.8 5.1 2.1 of which: isoC7 9.47 11.1 5.1 2.1 toluene 19.7 27.2 - - xylene 20.1 27.7 - - Total 100 100 100 100 RON 95.5 100.1 77.6 80.5 MY 85.8 89.1 74.5 77.8 (RON + MON) / 2 90.6 94.6 76.1 79.1

Example 2 :

The process described in Example 1 is reproduced, with the same apparatus, the same hydrogenation and isomerization catalysts, and the same operating conditions, except as regards the distillation column, including the regulation regulation of the bottom temperature is fixed at 188 ° C. Thus the overhead effluent from the distillation zone is practically free of cyclohexane and isoparaffins with 7 carbon atoms per molecule.

A residue and a distillate are collected at the bottom and at the head of the distillation column, respectively with a flow rate of 195.7 and 54.2 g / h, the compositions and octane numbers of which are presented in the third and fourth columns of Table 2. In the last column of the table are the composition and the octane numbers of the isomerate.

Compared to Example 1, the distillate has a much lower cyclohexane content and a very low content of isoparaffins with 7 carbon atoms per molecule. Its isomerization allows the average octane number to be raised by more than 10 points, practically without loss in the form of very volatile products (C ₃ ^- ). By mixing the isomerate with the residue from the distillation, a reconstituted gasoline is obtained which is almost free of benzene and olefins, having an average octane number equal to 90.8, ie substantially higher than that of the starting reformate, and without significant loss of yield. Table 2: compositions (% by weight) and octane numbers of the different streams for example 2 Reformate Residue Distillate Isomerate Hydrocarbons C6 ^- 26.4 6.1 99.9 99.9 of which: C3- - - - 0.08 olefins 0.19 - - - benzene 4.70 0.01 0.54 - cyclohexane 0.08 5.83 0.43 1.27 Hydrocarbons C7 + 73.6 93.9 0.18 0.1 of which: isoC7 9.47 12.1 0.18 0.1 toluene 19.7 25.2 - - xylene 20.1 25.6 - - Total 100 100 100 100 RON 95.5 98.5 72.5 83.3 MY 85.8 87.6 71.6 82.3 (RON + MON) / 2 90.6 93.1 72.1 82.8

Claims (31)

Process for the treatment of a charge, consisting mainly of hydrocarbons comprising at least 5 carbon atoms per molecule and comprising at least one unsaturated compound comprising at most six carbon atoms per molecule including benzene, such as: - said feed is treated in a distillation zone, comprising a depletion zone and a rectification zone, associated with a reaction hydrogenation zone, comprising at least one catalytic bed, in which the hydrogenation of at least part of the unsaturated compounds, comprising at most six carbon atoms per molecule and contained in the charge, in the presence of a hydrogenation catalyst and of a gas stream comprising hydrogen, the charge of the reaction zone being removed at the level of a sampling level and representing at least part of the liquid flowing in the rectification zone, the effluent from the reaction zone being at least partly reintroduced into the distillation zone, so as to ensure the continuity of distillation, and so as to finally leave at the head of the distillation zone an effluent very depleted in unsaturated compounds comprising at most six carbon atoms per mo lecule and at the bottom of the distillation zone an effluent depleted in unsaturated compounds comprising at most six carbon atoms per molecule, - At least part of the effluent, withdrawn at the head of the distillation zone in an isomerization zone, is treated, said part containing paraffins containing 5 and / or 6 carbon atoms per molecule in the presence of a catalyst isomerization, so as to obtain an isomerate. Process according to Claim 1, such that the distillation is carried out under a pressure of between 2 and 20 bar, with a reflux rate of between 1 and 10, the head temperature of the distillation zone being between 40 and 180 ° C and the distillation zone bottom temperature being between 120 and 280 ° C. Process according to either of Claims 1 and 2, in which the hydrogenation reaction zone is at least partially internal to the distillation zone. Process according to either of Claims 1 and 2, in which the hydrogenation reaction zone is at least partially external to the distillation zone. Process according to either of Claims 1 and 2, in which the hydrogenation zone is both partially incorporated in the rectification zone of the distillation zone and partially external to the distillation zone. Method according to one of claims 3 or 5 such that, for the part of the hydrogenation reaction internal to the distillation zone, the hydrogenation reaction is carried out at a temperature between 100 and 200 ° C, at a pressure between 2 and 20 bar, and the flow rate of hydrogen supplying the hydrogenation zone is between once and 10 times the flow rate corresponding to the stoichiometry of the hydrogenation reactions involved. Method according to one of claims 4 or 5 such that, for the part of the hydrogenation reaction external to the distillation zone, the pressure required for this hydrogenation step is between 1 and 60 bar, the temperature is between between 100 and 400 ° C, the space speed within the hydrogenation zone, calculated with respect to the catalyst, is generally between 1 and 50 h ^-1 (volume of charge per volume of catalyst and per hour), and the hydrogen flow rate corresponding to the stoichiometry of the hydrogenation reactions involved is between 0.5 and 10 times said stoichiometry. Method according to one of claims 3, 5 or 6 such that the hydrogenation catalyst is in contact with a descending liquid phase and with an ascending vapor phase, for any catalytic bed of the internal part of the hydrogenation zone. A method according to claim 8 such that the gas stream comprising hydrogen necessary for the hydrogenation zone is joined to the vapor phase, substantially at the inlet of at least one catalytic bed of the hydrogenation zone. Method according to one of claims 1 to 7 such that the flow of the liquid to be hydrogenated is co-current with the flow of the gas flow comprising hydrogen, for any catalytic bed of the internal part of the hydrogenation zone . Method according to one of claims 3, 5 or 6 such that the flow of the liquid to be hydrogenated is co-current with the flow of the gas stream comprising hydrogen and such that the distillation vapor is practically not in contact with the catalyst, for any catalytic bed in the internal part of the hydrogenation zone. Process according to Claim 11, such that the hydrogenation zone comprises at least one device for distributing liquid in any catalytic bed in said zone and at least one device for distributing the gas flow comprising hydrogen, for any catalytic bed in the hydrogenation zone internal to said zone. A method according to claim 12 such that the device for distributing the gas stream comprising hydrogen is arranged before the liquid dispensing device. A method according to claim 12 such that the device for distributing the gas flow comprising hydrogen is disposed at the level of the liquid distribution device. A method according to claim 12 such that the device for distributing the gas stream comprising hydrogen is arranged after the liquid dispensing device. Method according to one of claims 1 to 15 such that the bottom effluent from the distillation zone is mixed at least in part with the isomerization effluent. Method according to one of claims 1 to 16 such that the overhead effluent from the distillation zone is practically free of cyclohexane and isoparaffins with 7 carbon atoms per molecule. Method according to one of claims 1 to 17 such that the catalyst used in the hydrogenation zone comprises at least one metal chosen from the group formed by nickel and platinum. Method according to one of claims 1 to 18 such that the catalyst used in the hydrogenation zone comprises a support. Method according to one of claims 1 to 19 such that the isomerization catalyst comprises at least one metal from group VIII of the periodic table of the elements and a support comprising alumina. The method of claim 20 such that said catalyst further comprises at least one halogen. Method according to one of claims 20 or 21 such that the temperature is between 80 and 300 ° C, the partial pressure of hydrogen is between 0.1 and 70 bar, the space velocity is between 0.2 and 10 liters of liquid hydrocarbons per liter and catalyst per hour, and the molar ratio of hydrogen to hydrocarbons in the isomerate is greater than 0.06. Method according to one of claims 1 to 19 such that the isomerization catalyst comprises at least one metal from group VIII of the periodic table of the elements and a zeolite. The method of claim 23 such that said zeolite is selected from the group formed by mordenite and omega zeolite. Method according to one of claims 23 or 24 such that the temperature is between 200 and 300 ° C, the partial pressure of hydrogen is between 0.1 and 70 bar, the space speed is between 0.5 and 10 liters of liquid hydrocarbons per liter of catalyst per hour, and the molar ratio of hydrogen to hydrocarbon in the isomerate is between 0.07 and 15. Method according to one of claims 20 to 25 such that the metal of group VIII is chosen from the group formed by platinum, nickel and palladium. Method according to one of claims 1 to 26 such that the hydrogen, possibly in excess, which leaves at the head of the distillation zone can be recovered, then compressed and reused in the hydrogenation zone. Method according to one of claims 1 to 26 such that the hydrogen, possibly in excess, which leaves at the head of the distillation zone can be recovered, then compressed and reused in the isomerization zone. Method according to one of claims 1 to 26 such that hydrogen, possibly in excess, which leaves at the head of the distillation zone is recovered, then injected upstream of the compression stages associated with a catalytic reforming unit, in mixture with hydrogen from said unit. A method according to claim 29 such that said catalytic reforming unit operates at a pressure of less than 8 bar. Method according to one of claims 1 to 30, such that another cut comprising paraffins containing mainly 5 and / or 6 carbon atoms per molecule is also treated in the isomerization zone.

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