EP0781831A1 - Verfahren zur Erniedrigung des Gehaltes von Benzol und von leichten ungesättigten Verbindungen in Kohlenwasserstofffraktionen - Google Patents

Verfahren zur Erniedrigung des Gehaltes von Benzol und von leichten ungesättigten Verbindungen in Kohlenwasserstofffraktionen Download PDF

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EP0781831A1
EP0781831A1 EP96402910A EP96402910A EP0781831A1 EP 0781831 A1 EP0781831 A1 EP 0781831A1 EP 96402910 A EP96402910 A EP 96402910A EP 96402910 A EP96402910 A EP 96402910A EP 0781831 A1 EP0781831 A1 EP 0781831A1
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Prior art keywords
zone
hydrogenation
hydrogen
distillation
catalyst
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EP96402910A
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English (en)
French (fr)
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EP0781831B1 (de
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Christine Travers
Jean Cosyns
Charles Cameron
Jean-Luc Nocca
Francoise Montecot
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IFP Energies Nouvelles IFPEN
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IFP Energies Nouvelles IFPEN
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
    • C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
    • C10G65/04—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only refining steps
    • C10G65/08—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including only refining steps at least one step being a hydrogenation of the aromatic hydrocarbons
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline

Definitions

  • 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.
  • 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.
  • 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.
  • Benzene from a reformate can also be hydrogenated to cyclohexane.
  • 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.
  • ethers such as MTBE or ETBE
  • branched paraffinic hydrocarbons can be generated from the reformate itself, by isomerization of the linear paraffins.
  • the catalysts for isomerization of linear paraffins into branched paraffins are not inactive with respect to hydrocarbons of other chemical families.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • the external part of the hydrogenation zone is supplied by at least two sampling levels.
  • 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.
  • 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.
  • 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).
  • 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.
  • 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.
  • 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.
  • the excess hydrogen which leaves the top of the distillation zone is recovered, then compressed and reused in the isomerization zone.
  • 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.
  • hydrogen from catalytic reforming processes, methanation, P.S.A. (alternating pressure adsorption), electrochemical generation, steam cracking or steam reforming.
  • 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 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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 .
  • the device for distributing the gas flow comprising hydrogen is placed before the liquid dispensing device, and therefore before the catalytic bed.
  • 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.
  • 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.
  • 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.
  • 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 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.
  • 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.
  • 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 .
  • 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.
  • the catalyst used in the hydrogenation zone 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.
  • any other hydrogenation catalyst known to those skilled in the art can also be chosen.
  • the catalyst can advantageously contain at least one halogen in a proportion by weight relative to the catalyst of between 0.2 and 2%.
  • 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%.
  • 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.
  • 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%.
  • 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%.
  • 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.
  • it comprises at least one metal from group VIII of the periodic table of the elements and a support comprising alumina.
  • a support comprising alumina.
  • it also comprises at least one halogen, preferably chlorine.
  • 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.
  • 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.
  • 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 ⁇ zeolite.
  • a mordenite having a Si / Al (atomic) ratio between 5 and 50 and preferably between 5 and 30 is preferably used.
  • 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.
  • a catalyst based on an omega ⁇ or mazzite 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.
  • 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).
  • its DX crystallinity level 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.
  • 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.
  • the content by weight is between 0.05 and 1% and preferably between 0.1 and 0.6%.
  • 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.
  • 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.
  • 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.
  • the temperature is generally between 200 and 300 ° C and preferably between 230 and 280 ° C
  • 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.
  • FIG. 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.
  • FIG. 1 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.
  • 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.
  • 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.
  • the crude C 5 + reformate 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).
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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 100 100 100 100 100 100 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 1 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.
  • 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.
  • the distillate 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 - ).
  • 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 100 100 100 100 100 100 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

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  • Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
  • Vaporization, Distillation, Condensation, Sublimation, And Cold Traps (AREA)
EP96402910A 1995-12-27 1996-12-27 Verfahren zur Erniedrigung des Gehaltes von Benzol und von leichten ungesättigten Verbindungen in Kohlenwasserstofffraktionen Expired - Lifetime EP0781831B1 (de)

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FR9515529A FR2743081B1 (fr) 1995-12-27 1995-12-27 Procede de reduction selective de la teneur en benzene et en composes insatures legers d'une coupe d'hydrocarbures

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US6294492B1 (en) 1999-06-30 2001-09-25 Philips Petroleum Company Catalytic reforming catalyst activation
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MXPA06015023A (es) * 2006-12-19 2008-10-09 Mexicano Inst Petrol Aplicacion de material adsorbente microporoso de carbon, para reducir el contenido de benceno de corrientes de hidrocarburos.
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US9315741B2 (en) * 2008-09-08 2016-04-19 Catalytic Distillation Technologies Process for ultra low benzene reformate using catalytic distillation
US7910070B2 (en) * 2008-12-09 2011-03-22 Uop Llc Process for reducing benzene concentration in reformate
US20100145118A1 (en) * 2008-12-09 2010-06-10 Zimmerman Cynthia K Process for Reducing Benzene Concentration in Reformate
US8808533B2 (en) * 2010-04-23 2014-08-19 IFP Energies Nouvelles Process for selective reduction of the contents of benzene and light unsaturated compounds of different hydrocarbon fractions
EP2277980B1 (de) 2009-07-21 2018-08-08 IFP Energies nouvelles Verfahren zur selektiven reduzierung des benzolgehalts und des gehalts an leichten ungesättigten verbindungen von verschiedenen kohlenwasserstoffverschnitten
KR101835928B1 (ko) 2012-02-01 2018-03-07 사우디 아라비안 오일 컴퍼니 감소된 벤젠 가솔린을 생산하기 위한 촉매 개질 공정 및 시스템
CN104203388B (zh) * 2012-02-13 2018-01-02 反应性精馏技术有限公司 用于进行化学反应的反应性精馏塔
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DE69617892D1 (de) 2002-01-24
JPH09202886A (ja) 1997-08-05
CA2194085A1 (fr) 1997-06-28
FR2743081A1 (fr) 1997-07-04
EP0781831B1 (de) 2001-12-12
KR970033028A (ko) 1997-07-22
KR100447857B1 (ko) 2004-11-12
US6048450A (en) 2000-04-11
CA2194085C (fr) 2007-03-06
FR2743081B1 (fr) 1998-01-30
JP3806810B2 (ja) 2006-08-09
DE69617892T2 (de) 2002-04-25

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