Classifications

• • 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
• • 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
  • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
• • 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
  • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
  • C10G45/32—Selective hydrogenation of the diolefin or acetylene compounds
• • 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
  • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
  • C10G45/32—Selective hydrogenation of the diolefin or acetylene compounds
  • C10G45/34—Selective hydrogenation of the diolefin or acetylene compounds characterised by the catalyst used
  • C10G45/36—Selective hydrogenation of the diolefin or acetylene compounds characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof
  • C10G45/38—Selective hydrogenation of the diolefin or acetylene compounds characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof in combination with chromium, molybdenum or tungsten metals, or compounds thereof
• • 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
  • C10G65/06—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 selective hydrogenation of the diolefins
• • 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
  • C10G65/14—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural parallel stages only
• • 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
  • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
  • C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
• • 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
  • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
  • C10G2300/10—Feedstock materials
  • C10G2300/1037—Hydrocarbon fractions
  • C10G2300/104—Light gasoline having a boiling range of about 20 - 100 °C
• • 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
  • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
  • C10G2300/10—Feedstock materials
  • C10G2300/1037—Hydrocarbon fractions
  • C10G2300/1044—Heavy gasoline or naphtha having a boiling range of about 100 - 180 °C
• • 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
  • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
  • C10G2300/20—Characteristics of the feedstock or the products
  • C10G2300/201—Impurities
  • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
• • 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
  • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
  • C10G2300/40—Characteristics of the process deviating from typical ways of processing
  • C10G2300/4006—Temperature

Definitions

• the present invention relates to a process for the production of gasoline with low sulfur and mercaptan content.
• conversion gasolines and more particularly those obtained from catalytic cracking, which can represent 30 to 50% of the gasoline pool, have high mono-olefins and sulfur contents.
• FCC Fluid Catalytic Cracking according to English terminology
• gasolines obtained from catalytic cracking processes which will be called FCC (Fluid Catalytic Cracking according to English terminology) gasoline in the following. This can be translated as catalytic cracking in a fluidized bed).
• FCC gasolines therefore constitute the preferred feedstock for the process of the present invention.
• the one which has been very widely adopted consists in specifically treating the sulfur-rich gasoline bases by catalytic hydrodesulfurization processes in the presence of hydrogen.
• Traditional processes desulphurize gasoline in a non-selective manner by hydrogenating a large part of the mono-olefins, which generates a high loss in octane number and a high consumption of hydrogen.
• the most recent processes such as the Prime G + process (trademark), make it possible to desulfurize Cracking gasolines rich in olefins, while limiting the hydrogenation of monoolefins and consequently the loss of octane and the high consumption of hydrogen which results therefrom.
• the residual sulfur compounds generally present in desulphurized gasoline can be separated into two distinct families: the unconverted refractory sulfur compounds present in the feed on the one hand, and the sulfur compounds formed in the reactor by secondary reactions known as recombination.
• the majority compounds are the mercaptans resulting from the addition of the hhS formed in the reactor on the mono-olefins present in the feed.
• Mercaptans with the chemical formula R-SH, where R is an alkyl group are also called recombinant mercaptans. Their formation or decomposition obeys the thermodynamic equilibrium of the reaction between mono-olefins and hydrogen sulfide to form recombinant mercaptans. An example is illustrated according to the following reaction:
• Sulfur in the form of recombinant mercaptans generally represents between 20% and 80% by weight of the residual sulfur in desulfurized gasolines.
• An object of the present invention is to provide a process for treating a gasoline containing sulfur, part of which is in the form of mercaptans, which makes it possible to reducing the mercaptan content of said hydrocarbon fraction while limiting the loss of octane as much as possible.
• the first stage is generally intended to carry out a deep desulfurization of gasoline with a minimum of olefin saturation (and no aromatic loss) leading to a maximum retention of octane.
• the catalyst used is generally a CoMo type catalyst.
• new sulfur compounds are formed by recombination of the hhS resulting from desulfurization and olefins: the recombinant mercaptans.
• the second step generally has the role of minimizing the amount of recombinant mercaptans.
• the gasoline is then treated in a hydrodesulfurization reactor also called a finishing reactor with a catalyst generally based on nickel which exhibits practically no olefin hydrogenation activity and is capable of reducing the quantity of recombination mercaptans.
• the temperature is generally higher in the finishing reactor in order to thermodynamically promote the removal of the mercaptans.
• a furnace is therefore placed between the two reactors in order to be able to raise the temperature of the second reactor to a temperature higher than that of the first.
• the hydrogen used in the two stages is completely injected into the selective HDS reactor, the quantity of hydrogen entering into the finishing reactor being undergone and equal to the quantity injected into the first reactor minus the hydrogen consumed in this first reactor.
• the operating temperature is generally low in order to sufficiently desulfurize gasoline without generating a strong hydrogenation of the olefins.
• a reactor that is too cold can generate several problems, in particular a two-phase flow and no longer 100% gas, potentially inducing hydrodynamic problems or the impossibility of reaching a sufficiently high temperature in the finishing reactor to achieve a satisfactory conversion of the mercaptans from recombination, the heating power of the intermediate furnace being limited.
• a known solution from the prior art is then to simultaneously lower the ratio between the hydrogen flow rate and the feed rate to be treated, also called hereafter the H 2 / HC ratio, and to increase the temperature of the first reactor.
• the negative influence of the decrease in the H 2 / HC ratio on the reactions of hydrodesulfurization and hydrogenation of olefins is compensated by the increase in temperature.
• Increasing the temperature in the first reactor then makes it possible to adjust the temperature of the finishing reactor to a higher value.
• the induced drop in the H 2 / HC ratio in the finishing reactor has a negative effect on the thermodynamics of the reaction for eliminating the recombinant mercaptans, the partial pressures of H 2 S and of olefins being higher.
• the objective of the present invention is to remedy the drawbacks of the prior art by using, in a chain of two reactors without elimination of the hhS between the two stages, a higher H 2 / HC ratio in the finishing step than in the selective HDS step. This is achieved by injecting hydrogen (fresh or recycled) upstream of the finishing reactor.
• a higher H 2 / HC ratio in the finishing reactor in particular makes it possible to maintain a high temperature in the first reactor (and therefore also in the finishing reactor), while lowering the partial pressures of H2S and of olefins in the finishing reactor to optimize the conversion of the recombination mercaptans.
• the subject of the invention is a process for the treatment of a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following stages: a) contacting in at least one reactor, 'gasoline, hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase comprising a metal from group VIB and a metal from group VIII, at a temperature between 210 and 320 ° C, at a pressure between 1 and 4 MPa, with a space velocity between 1 and 10 fr 1 and a ratio between the hydrogen flow rate expressed in normal m 3 per hour and the feed rate to be treated expressed in m 3 per hour under standard conditions between 100 Nm 3 / m 3 and 600 Nm 3 / m 3 , so as to convert at least part of the sulfur compounds into H2S, b) at least part of the effluent is brought into contact in at least one reactor from step a) without removing the hhS formed, the hydrogen and u n hydrodesul
• Another advantage of the method according to the invention comes from the fact that it can easily be installed on existing units (remodeling or revamping according to English terminology).
• the ratio of the ratio between the hydrogen flow rate and the feed rate to be treated at the inlet of the reactor of step b) / ratio between the hydrogen flow rate and the feed rate to be treated at l input of the reactor of step a) is greater than or equal to 1.05.
• the ratio is between 1, 1 and 4.
• step c fresh hydrogen is injected in step c).
• the temperature of step b) is at least 5 ° C higher than the temperature of step a).
• the catalyst of step a) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight relative to the total weight of oxide catalyst of cobalt, in CoO form, between 0.1 and 10%, a content by weight relative to the total weight of molybdenum oxide catalyst, in M0O3 form, between 1 and 20%, a cobalt / molybdenum molar ratio of between between 0.1 and 0.8, a content by weight relative to the total weight of phosphorus oxide catalyst in P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface area between 30 and 180 m 2 / g.
• the catalyst of step b) consists of alumina and nickel, said catalyst containing a content by weight relative to the total weight of catalyst of nickel oxide, in NiO form, of between 5 and 20 %, said catalyst having a specific surface area between 30 and 180 m 2 / g.
• step c) of separating the effluent from step b) is carried out in a debutanizer or a stripping section.
• a gasoline distillation step is carried out so as to fractionate said gasoline into at least two light and heavy gasoline cuts and the heavy gasoline cut is treated in stages a), b). and c).
• the gasoline before step a) and before any possible distillation step, the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst in order to selectively hydrogenate the diolefins contained in said gasoline to olefins.
• the gasoline is catalytic cracked gasoline.
• step b) is carried out in at least two reactors in parallel.
• the H 2 / HC ratio of step b) is the same for each reactor in parallel.
• step b ′ carried out in parallel with step b)
• another part of the effluent from step a) is brought into contact in at least one reactor without eliminating the hS formed, hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, at a temperature between 280 and 400 ° C, at a pressure included between 0.5 and 5 MPa, with a space speed of between 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m 3 per hour and the feed rate to be treated expressed in m 3 per hour under conditions standards between 100 and 600 Nm 3 / m 3 , said temperature of step b ′) being higher than the temperature of step a).
• group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification.
• the metal content is measured by X-ray fluorescence.
• FIG. 1 illustrates an embodiment according to the invention.
• FIG. 2 illustrates another embodiment according to the invention.
• FIG. 3 illustrates another embodiment according to the invention. Detailed description of the invention Description of the load
• the process according to the invention makes it possible to treat any type of gasoline cut containing sulfur compounds and olefins, such as for example a cut obtained from a coking unit (coking according to the English terminology), visbreaking (visbreaking according to Anglo-Saxon terminology), steam cracking (steam cracking according to Anglo-Saxon terminology) or catalytic cracking (FCC, Fluid Catalytic Cracking according to Anglo-Saxon terminology).
• This essence can optionally be composed a significant fraction of gasoline from other production processes such as atmospheric distillation (gasoline resulting from direct distillation (or straight run gasoline according to the English terminology) or from conversion processes (coking gasoline or Said feed preferably consists of a gasoline cut obtained from a catalytic cracking unit.
• the feed is a gasoline cut containing sulfur compounds and olefins whose range of boiling points typically extends from the boiling points of hydrocarbons with 2 or 3 carbon atoms (C2 or C3) up to 260 ° C, preferably from the boiling points of hydrocarbons with 2 or 3 carbon atoms (C2 or C3) up to 220 ° C, more preferably from the boiling points of hydrocarbons with 5 carbon atoms up to at 220 ° C.
• the method according to the invention can also treat feeds having end points lower than those mentioned above, such as for example a C5-180 ° C cut.
• the sulfur content of gasoline cuts produced by catalytic cracking depends on the sulfur content of the feed treated by the FCC, the presence or not of a pretreatment of the feed from the FCC, as well as the end point of the feed. chopped off.
• the sulfur contents of an entire gasoline cut, especially those from the FCC are greater than 100 ppm by weight and most of the time greater than 500 ppm by weight.
• the sulfur contents are often greater than 1000 ppm by weight, they can even in certain cases reach values of the order of 4000 to 5000 ppm by weight.
• gasolines obtained from catalytic cracking units contain, on average, between 0.5% and 5% by weight of diolefins, between 20% and 50% by weight of olefins, between
• Hydrodesulfurization step a) is carried out to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H2S which is then removed in step c). Its implementation is particularly necessary when the feed to be desulfurized contains more than 100 ppm by weight of sulfur and more generally more than 50 ppm by weight of sulfur.
• Hydrodesulfurization step a) consists of bringing the gasoline to be treated into contact with hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization.
• step a) is carried out with the aim of carrying out a hydrodesulfurization selectively, that is to say with a degree of hydrogenation of the mono-olefins of less than 80%, preferably less than 70% and very preferably less than 60%.
• the temperature is generally between 210 and 320 ° C and preferably between 220 and 290 ° C. The temperature used must be sufficient to maintain the gasoline to be treated in the vapor phase in the reactor.
• the temperature of each reactor is generally higher by at least 5 ° C, preferably by at least 10 ° C and very preferably at least 30 ° C at the temperature of the reactor preceding it.
• the operating pressure of this step is generally between 1 and 4 MPa and preferably between 1, 5 and 3 MPa.
• the quantity of catalyst used in each reactor is generally such that the ratio between the gasoline flow rate to be treated expressed in m 3 per hour at standard conditions, per m 3 of catalyst (also called space speed) is between 1 and 10 h 1 and preferably between 2 and 8 h 1 .
• the hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m 3 per hour (Nm 3 / h) and the feed rate to be treated expressed in m 3 per hour at standard conditions (15 ° C , 0.1 MPa) is between 100 and 600 Nm 3 / m 3 , preferably between 200 and 500 Nm 3 / m 3 .
• Normal m 3 is understood to mean the quantity of gas in a volume of 1 m 3 at 0 ° C. and 0.1 MPa.
• the hydrogen required for this step can be fresh hydrogen or recycled hydrogen, preferably free of hhS, or a mixture of fresh hydrogen and recycled hydrogen. Preferably, fresh hydrogen will be used.
• the degree of desulfurization of step a), which depends on the sulfur content of the feed to be treated, is generally greater than 50% and preferably greater than 70% so that the product resulting from step a) contains less than 100 ppm by weight of sulfur and preferably less than 50 ppm by weight of sulfur.
• the catalyst used in step a) must exhibit good selectivity with respect to the hydrodesulfurization reactions relative to the hydrogenation reaction of olefins.
• the hydrodesulfurization catalyst of step a) comprises an oxide support and an active phase comprising a metal from group VI B and a metal from group VIII and optionally phosphorus and / or an organic compound as described below. .
• the group VI B metal present in the active phase of the catalyst is preferably chosen from molybdenum and tungsten.
• the group VIII metal present in the phase active catalyst is preferably chosen from cobalt, nickel and a mixture of these two elements.
• the active phase of the catalyst is preferably chosen from the group formed by the combination of the elements nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum and very preferably the active phase consists of cobalt and molybdenum.
• the group VIII metal content is between 0.1 and 10% by weight of oxide of the group VIII metal relative to the total weight of the catalyst, preferably between 0.6 and 8% by weight, preferably between 2 and 7% by weight, very preferably between 2 and 6% by weight and even more preferably between 2.5 and 6% by weight.
• the group VI B metal content is between 1 and 20% by weight of oxide of the metal of group VI B relative to the total weight of the catalyst, preferably between 2 and 18% by weight, very preferably between 3 and 16% by weight.
• the group VIII metal to group VI B metal molar ratio of the catalyst is generally between 0.1 and 0.8, preferably between 0.2 and 0.6.
• the catalyst has a density of metal from group VI B, expressed as the number of atoms of said metal per unit area of the catalyst, which is between 0.5 and 30 atoms of metal from group VI B per nm 2 of catalyst, preferably between 2 and 25, even more preferably between 3 and 15.
• the density of metal from group VI B expressed in number of atoms of metal from group
• VI B per unit area of the catalyst (number of atoms of metal from group VI B per nm 2 of catalyst) is calculated for example from the following relationship: (CCN a )
• N A Avogadro number equal to 6,022.10 23 ;
• S Specific surface area of the catalyst (m 2 / g), measured according to standard ASTM D3663;
• M M Molar mass of the metal of group VIB (for example 95.94 g / mol for molybdenum).
• the density d (Mo) is equal to :
• the catalyst may also have a phosphorus content generally between 0.3 and 10% by weight of P2O5 relative to the total weight of catalyst, preferably between 0.5 and 5% by weight, very preferably between 1 and 3% by weight.
• the phosphorus present in the catalyst is combined with the metal of group VIB and optionally also with the metal of group VIII in the form of heteropolyanions.
• the phosphorus / (group VIB metal) molar ratio is generally between 0.1 and 0.7, preferably between 0.2 and 0.6, when phosphorus is present.
• the catalyst is characterized by a specific surface area of between 5 and 400 m 2 / g, preferably between 10 and 250 m 2 / g, preferably between 20 and 200 m 2 / g, very preferably. between 30 and 180 m 2 / g.
• the specific surface is determined in the present invention by the BET method according to the ASTM D3663 standard, as described in the work Rouquerol F .; Rouquerol J .; Singh K. “Adsorption by Powders & Porous So / ids; Princip / e, methodology and applications ”, Academy Press, 1999, for example by means of an Autopore III TM model apparatus of the brand Microméritics TM.
• the total pore volume of the catalyst is generally between 0.4 cm 3 / g and 1.3 cm 3 / g, preferably between 0.6 cm 3 / g and 1.1 cm 3 / g.
• the total pore volume is measured by mercury porosimetry according to standard ASTM D4284 with a wetting angle of 140 °, as described in the same book.
• the packed fill density (DRT) of the catalyst is generally between 0.4 and 0.7 g / mL, preferably between 0.45 and 0.69 g / mL.
• the DRT measurement consists in introducing the catalyst into a test tube whose volume has been determined beforehand and then, by vibration, in compacting it until a constant volume is obtained.
• the bulk density of the packed product is calculated by comparing the introduced mass and the occupied volume after packing.
• the hydrodesulphurization catalyst before sulphurization, has an average pore diameter greater than 20 nm, preferably greater than 25 nm, or even 30 nm and often between 20 and 140 nm, preferably between 20 and 100 nm, and very preferably between 25 and 80 nm.
• the pore diameter is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140 °.
• the catalyst can be in the form of extrudates of small diameter, cylindrical or multilobed (trilobes, quatrefoils, etc.), or spheres.
• the catalyst oxide support is usually a porous solid chosen from the group consisting of: aluminas, silica, alumina silicas or even titanium or magnesium oxides used alone or as a mixture with alumina or silica alumina . It is preferably chosen from the group consisting of silica, the family of transition aluminas and alumina silicas, very preferably, the oxide support consists essentially of alumina, that is to say that it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight of alumina. It preferably consists solely of alumina.
• the oxide support of the catalyst is a “high temperature” alumina, that is to say which contains theta, delta, kappa or alpha phase aluminas, alone or as a mixture, and an amount of less than 20. % of gamma, chi or eta phase alumina.
• the catalyst can also further comprise at least one organic compound containing oxygen and / or nitrogen and / or sulfur before sulfurization. Such additives are described below.
• a very preferred embodiment of the invention corresponds to the implementation for step a) of a catalyst comprising alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight relative to the total weight of cobalt oxide catalyst, in CoO form, of between 0.1 and 10%, a content by weight relative to the total weight of molybdenum oxide catalyst, in M0O3 form, between 1 and 20%, a cobalt / molybdenum molar ratio of between 0.1 and 0.8, a content by weight relative to total weight of phosphorus oxide catalyst in P2O5 form between 0.3 and
• the active phase consists of cobalt and molybdenum. According to another embodiment, the active phase consists of cobalt, molybdenum and phosphorus.
• step b Description of the finishing hydrodesulfurization step (step b)
• hydrodesulfurization step a a large part of the sulfur compounds are converted into H2S.
• the remaining sulfur compounds are essentially refractory sulfur compounds and the recombinant mercaptans resulting from the addition of the hhS formed in step a) on the mono-olefins present in the feed.
• finishing hydrodesulfurization step b) is mainly carried out to decompose, at least in part, the recombinant mercaptans into olefins and H2S. Step b) also hydrodesulfurize the more refractory sulfur compounds.
• Step b) is carried out using a higher H2 / HC ratio and temperature than in step a) and in the presence of a particular catalyst.
• Step b) consists in bringing at least part of the effluent from step a) into contact with hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization .
• Hydrodesulfurization step b) is carried out without significant hydrogenation of the olefins.
• the degree of hydrogenation of the olefins of the catalyst of hydrodesulfurization step b) is generally less than 5% and even more generally less than 2%.
• the temperature of this step b) is generally between 280 and 400 ° C, more preferably between 300 and 380 ° C, and very preferably between 310 and 370 ° C.
• the temperature of this step b) is generally at least 5 ° C, preferably at least 10 ° C and very preferably at least 30 ° C above the temperature of step a).
• the operating pressure for this step is generally between 0.5 and 5 MPa and preferably between 1 and 3 MPa.
• the quantity of catalyst used in each reactor is generally such that the ratio between the gasoline flow rate to be treated expressed in m 3 per hour at standard conditions, per m 3 of catalyst (also called space speed) is between 1 and 10 h 1 and preferably between 2 and 8 h 1 .
• the ratio between the hydrogen flow rate and the feed flow rate to be treated also called the H2 / HC ratio of step b) is greater than the H2 / HC ratio of step a).
• the ratio between the hydrogen flow rate and the feed rate to be treated is meant the ratio at the inlet to the reactor of the stage concerned.
• the adjustment factor F is greater than or equal to 1.05, preferably greater than 1.1 and preferably between 1, 1 and 6, preferably between 1, 2 and 4, and preferably between 1, 2 and 2.
• step b In order to achieve such a H2 / HC ratio in step b), a supply of hydrogen is necessary. According to a preferred embodiment, fresh hydrogen is injected in step b).
• step b) recycled hydrogen preferably free of F ⁇ S beforehand.
• the recycled hydrogen can come from separation step c). It is also possible to inject a mixture of fresh and recycled hydrogen.
• step b) Some of the hydrogen present in step b) comes from step a) (hydrogen not consumed by the reactions that take place in step a)).
• the quantity of hydrogen injected only in step b) can be adjusted during the cycle, the deactivation of the catalyst of the first step a) being able to be compensated by a gradual increase in the H 2 / HC ratio in this reactor.
• step b) can be carried out in a plurality of reactors in parallel in order to minimize the size of said reactors and the gas surface velocity within said reactors.
• the catalyst of step b) is of a different nature and / or composition from that used in step a).
• the catalyst of step b) is in particular a very selective hydrodesulfurization catalyst: it makes it possible to hydrodesulfurize without hydrogenating the olefins and thus maintaining the octane number.
• the catalyst which may be suitable for this step b) of the process according to the invention is a catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, and of preferably chosen from the group formed by nickel, cobalt, iron. These metals can be used alone or in combination.
• the active phase consists of a metal from group VIII, preferably nickel.
• the active phase consists of nickel.
• the group VIII metal content is between 1 and 60% by weight of oxide of the group VIII metal relative to the total weight of the catalyst, preferably between 5 and 30% by weight, very preferably between 5 and 20. % weight.
• the catalyst is characterized by a specific surface area of between 5 and 400 m 2 / g, preferably between 10 and 250 m 2 / g, preferably between 20 and 200 m 2 / g, very preferably. between 30 and 180 m 2 / g.
• the specific surface is determined in the present invention by the BET method according to the ASTM D3663 standard, as described in the work Rouquerol F .; Rouquerol J .; Singh K. “Adsorption by Powders & Porous So / ids; Princip / e, methodology and applications ”, Academy Press, 1999, for example by means of an Autopore III TM model apparatus of the brand Microméritics TM.
• the pore volume of the catalyst is generally between 0.4 cm 3 / g and 1.3 cm 3 / g, preferably between 0.6 cm 3 / g and 1.1 cm 3 / g.
• the total pore volume is measured by mercury porosimetry according to standard ASTM D4284 with a wetting angle of 140 °, as described in the same book.
• the packed fill density (DRT) of the catalyst is generally between 0.4 and 0.7 g / mL, preferably between 0.45 and 0.69 g / mL.
• the DRT measurement consists in introducing the catalyst into a test tube, the volume of which has been determined beforehand, then, by vibration, in compacting it until a constant volume is obtained.
• the bulk density of the packed product is calculated by comparing the introduced mass and the occupied volume after packing.
• the catalyst of step b), before sulfurization has an average pore diameter greater than 20 nm, preferably greater than 25 nm, or even 30 nm and often between 20 and 140 nm, preferably between 20 and 100 nm, and very preferably between 25 and 80 nm.
• the pore diameter is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140 °.
• the catalyst can be in the form of extrudates of small diameter, cylindrical or multilobed (trilobes, quatrefoils, etc.), or spheres.
• the catalyst oxide support is usually a porous solid chosen from the group consisting of: aluminas, silica, alumina silicas or even titanium or magnesium oxides used alone or as a mixture with alumina or silica alumina . It is preferably chosen from the group consisting of silica, the family of transition aluminas and alumina silicas, very preferably, the oxide support consists essentially of alumina, that is to say that it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight of alumina. It preferably consists solely of alumina.
• the oxide support of the catalyst is a “high temperature” alumina, that is to say which contains theta, delta, kappa or alpha phase aluminas, alone or as a mixture, and an amount of less than 20. % of gamma, chi or eta phase alumina.
• a very preferred embodiment of the invention corresponds to the implementation for step b) of a catalyst consisting of alumina and nickel, said catalyst containing a content by weight relative to the total weight of catalyst of nickel oxide, in the form
• NiO between 5 and 20%, said catalyst having a specific surface area between 30 and
• the catalyst of the hydrodesulfurization step b) is characterized by a hydrodesulfurization catalytic activity generally between 1% and 90%, preferably between 1% and 70%, and very preferably between 1% and 50% of the catalytic activity of the catalyst of hydrodesulfurization step a).
• the rate of removal of the mercaptans from step b) is generally greater than 50% and preferably greater than 70% so that the product from step b) contains less than 10 ppm sulfur, and preferably less than 5 ppm sulfur from the recombinant mercaptans relative to the total weight of the feed.
• the degree of hydrogenation of the olefins of the catalyst of hydrodesulfurization step b) is generally less than 5% and even more generally less than 2%.
• the hydrodesulfurization steps a) and b) can be carried out in at least two different reactors.
• steps a) and b) are carried out using two reactors, the latter two are placed in series, the second reactor fully treating the effluent at the outlet of the first reactor (without separation of the liquid and the gas between the first and the second reactor) and by adding a flow of hydrogen between the two reactors such that the H2 / HC ratio at the inlet of step b) is greater than the H2 / HC ratio at the inlet of step a).
• the finishing step b) can be carried out in at least two reactors placed in parallel at the outlet of step a), without separation of the liquid and the gas at the outlet of said step a) and with an addition of hydrogen in each of the reactors of step b).
• step b) is carried out with two reactors.
• a hydrogen flow rate is added in each of the reactors so as to have a hh / HC ratio at the inlet of step b) greater than the H2 / HC ratio at the inlet of step a) as defined with the adjustment factor F.
• the reactors of step b) can be of equal or different volume.
• the hydrogen entering the finishing stage b) consists on the one hand of the hydrogen not consumed by the reactions which take place in the hydrodesulfurization stage a) and on the other hand of an addition of hydrogen (fresh and / or recycled, preferably free of FhS).
• the addition of hydrogen is preferably carried out at the outlet of step a) but upstream of the separation of the feed to the reactors in parallel with step b).
• the H 2 / HC ratio at the inlet of step b) is thus the same for each reactor in parallel with step b).
• the H 2 / HC ratio at the input of step b) is different for each reactor in parallel with step b), but greater than the H 2 / HC ratio of step a).
• the operating conditions according to this embodiment are the operating conditions described for step b) with a single reactor.
• the temperature of the reactors in parallel with step b) may or may not be the same.
• the temperature of the reactors of step b) is identical in the two reactors in parallel, which makes it possible to use a single furnace for heating the effluent of step a).
• a finishing step b ′) can be carried out in parallel with step b), step b) being carried out with an addition of hydrogen and step b ′) being carried out without addition of hydrogen, the two stages b) and b ') being carried out at temperatures higher than that of stage a).
• the quantity of hydrogen entering this step b ′) then being subjected to and equal to the quantity injected in step a) minus the hydrogen consumed in step a).
• stage b) Part of the effluent from step a) is thus subjected to stage b) carried out with a high H2 / HC ratio (by injecting hydrogen) while the other part of the effluent from stage a) is subjected in parallel to stage b ') without injection of additional hydrogen.
• all of the effluent from step a) is sent to steps b) and b ') (without separation of the liquid and gas between step a) and steps b). and b ') carried out in parallel).
• step b ') is carried out by bringing into contact, in at least one reactor, part of the effluent from step a) without removing the H2S formed, hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, at a temperature between 280 and 400 ° C, at a pressure between 0.5 and 5 MPa, with a space velocity between 1 and 10 fr 1 and a ratio between the hydrogen flow rate expressed in normal m 3 per hour and the feed rate to be treated expressed in m 3 per hour at standard conditions between 100 and 600 Nm 3 / m 3 , said temperature of step b ') being higher than the temperature of step a).
• the temperature of this step b ′) is generally between 280 and 400 ° C, more preferably between 300 and 380 ° C, and very preferably between 310 and 370 ° C.
• the temperature of this step b ') is generally at least 5 ° C, preferably at least 10 ° C and very preferably at least 30 ° C above the mean operating temperature of the machine. step a).
• step b ′ may or may not be the same as the temperature of step b).
• the operating pressure of this stage b ') is generally between 0.5 and 5 MPa and preferably between 1 and 3 MPa.
• the quantity of catalyst used in each reactor is generally such that the ratio between the gasoline flow rate to be treated expressed in m 3 per hour at standard conditions, per m 3 of catalyst (also called space speed) is between 1 and 10 h 1 and preferably between 2 and 8 h 1 .
• the hydrogen flow is undergone and equal to the amount injected in step a) minus the hydrogen consumed in step a).
• the hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m 3 per hour (Nm 3 / h) and the feed rate to be treated expressed in m 3 per hour at standard conditions (15 ° C , 0.1 MPa) is between 100 and 600 Nm 3 / m 3 , preferably between 200 and 500 Nm 3 / m 3 .
• the part of the effluent from step a) sent to step b) represents between 10 and 90% by volume, preferably between 20 and 80% by volume of the effluent from the step a).
• the part of the effluent from step a) sent to step b ’) corresponds to the effluent from step a) minus the effluent sent to step b).
• the part of the effluent from step a) sent to step b) is larger than the part of the effluent from step a) sent to step b ’).
• the catalyst of stage b ′) is a catalyst such as the catalyst described for stage b) of hydrodesulfurization.
• the catalyst of step b ′) can be identical or different from the catalyst of step b).
• a very preferred embodiment of the invention corresponds to the implementation for step b ′) of a catalyst consisting of alumina and nickel, said catalyst containing a content by weight relative to the total weight of catalyst d 'nickel oxide, in NiO form, between 5 and 20%, said catalyst having a specific surface area between 30 and 180 m 2 / g.
• step c) of the method is carried out with a step of separating the hhS formed and present in the effluent from step b).
• This step is carried out in order to separate the excess hydrogen as well as the HSH formed during steps a) and b). Any method known to those skilled in the art can be considered.
• the effluent from step b) is cooled to a temperature generally less than 80 ° C and preferably less than 60 ° C in order to condense the hydrocarbons.
• the gas and liquid phases are then separated in a separation flask.
• the liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved h ⁇ S is sent to a stabilization column or debutanizer. This column separates an overhead cut essentially consisting of residual hhS and hydrocarbon compounds having a boiling point less than or equal to that of butane and a bottom cut free of hhS, called stabilized gasoline, containing compounds having a boiling point higher than that of butane.
• the liquid fraction resulting from the effluent from step b) and which contains desulfurized gasoline as well as a fraction of dissolved hhS is sent to a section stripping, while the gaseous fraction consisting mainly of hydrogen and H2S is sent to a purification section.
• the stripping can be carried out by heating the hydrocarbon fraction alone or with an injection of hydrogen or water vapor, in a distillation column in order to to extract, at the head, the light compounds which were entrained by dissolution in the liquid fraction as well as the dissolved residual hhS.
• the temperature of the stripped gasoline recovered at the bottom of the column is generally between 120 ° C and 250 ° C.
• the separation step c) is carried out in a stabilization column or debutanizer.
• a stabilization column allows the HSH to be separated more efficiently than a stripping section.
• step b ′ When step b ′) is carried out in parallel with step b), the hhS formed and present in the effluent from step b ’) is separated in the same way.
• the effluent from step b ′) is introduced after cooling, mixed or not in the same separation flask as the effluent from step b), then into the same stabilization column or in the same stripping section.
• the effluent from step b ′) is introduced after cooling into a separation flask, the effluent from step b) is introduced into another separation flask, then the The resulting liquid fractions are introduced into the same stabilization column or the same stripping section.
• step b) When step b) is carried out in several reactors in parallel, the hhS formed and present in the effluent from each reactor of step b) is separated in the same way.
• each effluent from the reactors of step b) is introduced after cooling, mixed or not in the same separation flask, then in the same stabilization column or in the same stripping section.
• each effluent from step b) is introduced after cooling into a dedicated separation flask, then the fractions resulting liquids are introduced into the same stabilization column or the same stripping section.
• Step c) is preferably carried out so that the sulfur in the form of hhS remaining in the effluent from step b) represents less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.
• the hydrodesulfurization step b) or b ') respectively and the step c) of separation of the H ⁇ S when the hydrodesulfurization and the separation are carried out in parallel, without using the same means separation, can be carried out simultaneously by means of a catalytic column equipped with at least one catalytic bed containing the hydrodesulfurization catalyst.
• the catalytic distillation column has two beds of hydrodesulfurization catalyst and the effluent from step b) or b ') is sent to the column between the two catalyst beds.
• the preparation of the catalysts and of the sulfurization is known and generally comprises a stage of impregnation of metals of group VIII and of group VIB when it is present, and optionally phosphorus and / or organic compound on the oxide support, followed by drying, then by optional calcination making it possible to obtain the active phase in their oxide forms.
• the catalysts are generally subjected to sulfurization in order to form the active species as described below.
• the impregnation step can be carried out either by slurry impregnation, or by excess impregnation, or by dry impregnation, or by any other means known to those skilled in the art.
• the impregnation solution is chosen so as to be able to dissolve the metal precursors in the desired concentrations.
• the sources of molybdenum use may be made of oxides and hydroxides, molybdic acids and their salts, in particular ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PM012O40), and their salts, and optionally silicomolybdic acid (H4SNVI012O40) and its salts.
• the sources of molybdenum can also be any heteropolycompound such as Keggin, Lacunar Keggin, substituted Keggin, Dawson, Anderson, Strandberg, for example. Molybdenum trioxide and heteropolycompounds of Keggin, lacunar Keggin, substituted Keggin and Strandberg type are preferably used.
• the tungsten precursors which can be used are also well known to those skilled in the art.
• oxides and hydroxides, tungstic acids and their salts in particular ammonium salts such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and theirs. salts, and optionally silicotungstic acid (H4S1W12O40) and its salts.
• the sources of tungsten can also be any heteropolycompound such as Keggin, Lacunar Keggin, substituted Keggin, Dawson, for example.
• Oxides and ammonium salts, such as ammonium metatungstate or heteropolyanions of the Keggin, lacunar Keggin or substituted Keggin type are preferably used.
• cobalt precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example example. Cobalt hydroxide and cobalt carbonate are preferably used.
• the nickel precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example.
• the preferred phosphorus precursor is orthophosphoric acid H 3 PO 4 , but its salts and esters such as ammonium phosphates are also suitable.
• the phosphorus can also be introduced at the same time as the element (s) of group VIB in the form of heteropolyanions of Keggin, lacunar Keggin, substituted Keggin or of Strandberg type.
• the catalyst is generally subjected to a drying step at a temperature below 200 ° C, advantageously between 50 ° C and 180 ° C, preferably between 70 ° C and 150 ° C, of very preferably between 75 ° C and 130 ° C.
• the drying step is preferably carried out under an inert atmosphere or under an atmosphere containing oxygen.
• the drying step can be carried out by any technique known to those skilled in the art. It is advantageously carried out at atmospheric pressure or at reduced pressure. Preferably, this step is carried out at atmospheric pressure. It is advantageously carried out in a crossed bed using air or any other hot gas.
• the gas used is either air or an inert gas such as argon or nitrogen.
• the drying is carried out in a crossed bed in the presence of nitrogen and / or air.
• the drying step has a duration of between 5 minutes and 15 hours, preferably between 30 minutes and 12 hours.
• the catalyst has not undergone calcination during its preparation, that is to say that the impregnated catalytic precursor has not been subjected to a heat treatment step at a higher temperature. at 200 ° C under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.
• the catalyst has undergone a calcination step during its preparation, that is to say that the impregnated catalytic precursor has been subjected to a heat treatment step at a temperature between 250 and 1000 ° C and preferably between 200 and 750 ° C, for a period typically between 15 minutes and 10 hours, under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.
• the catalysts of the process according to the invention Before coming into contact with the feed to be treated in a gasoline hydrodesulfurization process, the catalysts of the process according to the invention generally undergo a sulfurization step.
• the sulphurization is preferably carried out in a sulforreducing medium, that is to say in the presence of hhS and hydrogen, in order to convert the metal oxides into sulphides such as, for example, MOS2, CogSs or N13S2.
• Sulfurization is carried out by injecting onto the catalyst a stream containing H 2 S and hydrogen, or else a sulfur compound capable of decomposing into H 2 S in the presence of the catalyst and of hydrogen.
• Polysulfides such as dimethyldisulfide (DM DS) are precursors of hhS commonly used to sulfide catalysts. Sulfur can also come from the charge. The temperature is adjusted so that the hhS reacts with the metal oxides to form metal sulphides. This sulfurization can be carried out in situ or
• the degree of sulfurization of the metals constituting the catalysts is at least equal to 60%, preferably at least equal to 80%.
• the sulfur content in the sulfurized catalyst is measured by elemental analysis according to ASTM D5373.
• a metal is considered to be sulphurized when the overall sulphurization rate defined by the molar ratio between the sulfur (S) present on the catalyst and said metal is at least equal to 60% of the theoretical molar ratio corresponding to the total sulphurization of the (s) metal (s) considered.
• the overall sulfurization rate is defined by the following equation:
• (S / metal) catalyst is the molar ratio between sulfur (S) and metal present on the catalyst
• Theoretical (S / metal) is the molar ratio between sulfur and metal corresponding to the total sulphurization of the metal to sulphide.
• the molar ratio between the S present on the catalyst and all the metals must also be at least equal to 60% of the theoretical molar ratio corresponding to the total sulphurization of each metal to sulphide, the calculation being carried out in proportion to the relative mole fractions of each metal.
• a step of distillation of the gasoline to be treated is carried out in order to separate two cuts (or fractions), namely a light cut and a heavy cut, and the heavy cut is treated according to the process of the invention.
• the heavy cut is treated by the method according to the invention.
• This first variant has the advantage of not hydrotreating the light cut which is rich in olefins and generally low in sulfur, which makes it possible to limit the loss of octane by hydrogenation of the olefins contained in the light cut.
• the light cut has a boiling temperature range below
• the gasoline to be treated is subjected before the hydrodesulfurization process according to the invention to a preliminary step consisting of a selective hydrogenation of the diolefins present in the feed, as described in patent application EP 1077247.
• the gasoline to be treated is treated beforehand in the presence of hydrogen and a selective hydrogenation catalyst so as to at least partially hydrogenate the diolefins and carry out a heavier reaction of part of the light mercaptan compounds (SHR) present. in the thioethers feed, by reaction with olefins.
• SHR light mercaptan compounds
• the gasoline to be treated is sent to a selective hydrogenation catalytic reactor containing at least one fixed or moving bed of catalyst for the selective hydrogenation of the diolefins and the weighting of the light mercaptans.
• the reaction for the selective hydrogenation of the diolefins and for the weighting down of the light mercaptans is preferably carried out on a sulfurized catalyst comprising at least one element from group VIII and optionally at least one element from group VI B and an oxide support.
• the element of group VIII is preferably chosen from nickel and cobalt and in particular nickel.
• the element of group VI B when it is present, is preferably chosen from molybdenum and tungsten and very preferably molybdenum.
• the oxide support of the catalyst is preferably chosen from alumina, nickel aluminate, silica, silicon carbide, or a mixture of these oxides.
• Alumina is preferably used and even more preferably high purity alumina is used.
• the selective hydrogenation catalyst contains nickel with a content by weight of nickel oxide, in NiO form, between 1 and 12%, and molybdenum with a content by weight of molybdenum oxide, in the form M0O3, between 1% and 18% and a nickel / molybdenum molar ratio of between 0.3 and 2.5, the metals being deposited on a support consisting of alumina.
• the degree of sulfurization of the metals constituting the catalyst is preferably greater than 60%.
• the gasoline is contacted with the catalyst at a temperature between 50 and 250 ° C, and preferably between 80 and 220 ° C, and even more preferably between 90 and 200 ° C, with a liquid space velocity (LHSV) of between 0.5 fr 1 and 20 h 1 , the unit of the liquid space speed being the liter of feed per liter of catalyst and per hour (L / L / h).
• LHSV liquid space velocity
• the pressure is between 0.4 and 5 MPa, preferably between 0.6 and 4 MPa and even more preferably between 1 and 3 MPa.
• the optional selective hydrogenation step is typically carried out with a ratio between the hydrogen flow rate expressed in normal m 3 per hour and the feed rate to be treated expressed in m 3 per hour at standard conditions of between 2 and 100 Nm 3 / m 3 , preferably between 3 and 30 Nm 3 / m 3 .
• the diolefin content determined by means of the maleic anhydride number (MAV or Maleic Anhydride Value according to the English terminology), according to the UOP 326 method, is generally reduced to less than 6 mg maleic anhydride / g, or even less than 4 mg AM / g and more preferably less than 2 mg AM / g. In some cases, it can be obtained less than 1 mg AM / g.
• MAV Maleic Anhydride number
• the selectively hydrogenated gasoline is then distilled into at least two cuts, a light cut and a heavy cut and optionally an intermediate cut.
• the heavy cut is treated according to the method of invention.
• the intermediate and heavy cuts can be treated separately by the process according to the invention.
• the gasoline to be treated and hydrogen are sent via line 1 via line 3 to a hydrodesulphurization unit 2 of step a).
• the gasoline treated is generally a cracked gasoline, preferably a catalytic cracked gasoline.
• Gasoline is characterized by a boiling point typically ranging between 30 ° C and 220 ° C.
• the hydrodesulfurization unit 2 of step a) is for example a reactor containing a supported hydrodesulfurization catalyst based on a metal from group VIII and VIB in a fixed bed or in a fluidized bed, preferably a reactor is used. in a fixed bed.
• the reactor is operated under operating conditions and in the presence of a hydrodesulfurization catalyst, as described above to decompose the sulfur compounds and form hydrogen sulfide (H2S).
• a hydrodesulfurization catalyst as described above to decompose the sulfur compounds and form hydrogen sulfide (H2S).
• H2S hydrogen sulfide
• step a recombinant mercaptans are formed by addition of hSs formed on the olefins.
• the effluent from the hydrodesulfurization unit 2 is then introduced into the hydrodesulfurization unit 5 of step b) via line 4 without removing the hSH formed.
• the hydrodesulfurization unit 5 is for example a reactor containing a catalyst hydrodesulfurization in a fixed bed or in a fluidized bed, preferably a fixed bed reactor is used.
• Unit 5 is operated at a higher temperature than unit 2 and in the presence of a particular catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII.
• Unit 5 is operated with a higher H2 / HC ratio than that of step a) in order to at least partially decompose the recombinant mercaptans into olefins and H2S by reduction of the pphhS.
• a supply of hydrogen is made via line 6. It also makes it possible to hydrodesulphurize, at least in part, the most refractory sulfur compounds.
• An effluent (gasoline) containing hSH is withdrawn from said hydrodesulphurization reactor 5 via line 7.
• step c The effluent then undergoes a step of removing the hhS (step c) which consists, in the embodiment of FIG. 1, in treating the effluent by condensation by introducing the effluent from step b) via line 7 in a separation flask 8 in order to withdraw a gas phase containing hhS and hydrogen via line 9 and a liquid fraction.
• the liquid fraction which contains the desulfurized gasoline as well as a fraction of dissolved hhS is sent via line 10 to a stabilization column or debutaniser 11 in order to separate at the top of the column via line 12 a flow containing C4 hydrocarbons and residual hhS and at the bottom via line 13 of the column a so-called stabilized gasoline containing compounds having a boiling point higher than that of butane.
• Figure 2 shows a second embodiment based on that of Figure 1 and which is distinguished by the presence of a finishing step b ’) without injection of hydrogen in parallel with step b).
• the gasoline to be treated is sent via line 1 and hydrogen via line 3 to a hydrodesulphurization unit 2 in step a).
• the hydrodesulfurization unit 15 is for example a reactor containing a hydrodesulfurization catalyst in a fixed bed or in a fluidized bed, preferably a fixed bed reactor is used. Unit 15 is operated at a higher temperature than unit 2 and in the presence of a particular catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII. No hydrogen supply is made in unit 15.
• step d An effluent (gasoline) containing hSH is withdrawn from said hydrodesulfurization reactor 15 via line 16.
• the effluent then undergoes a step of removing the hhS (step d) which consists, in the embodiment of FIG. 2, in treating the effluent by condensation by introducing the effluent from step b ') via line 16 in a separation flask 17 in order to withdraw a gas phase containing hhS and hydrogen via line 18 and a liquid fraction.
• the liquid fraction which contains desulphurized gasoline as well as a fraction of dissolved hhS is sent via line 19 to the stabilization column or debutanizer 11 in order to separate at the top of the column via line 12 a stream containing hydrocarbons in C4 and residual hhS and at the bottom via line 13 of the column a so-called stabilized gasoline containing the compounds having a boiling point higher than that of butane.
• FIG. 3 represents a third embodiment based on that of FIG. 2 and which is differentiated by the addition of hydrogen.
• the addition of hydrogen (6) is carried out at the outlet of step a) but upstream of the separation of the feed to the reactors in parallel with step b).
• the H2 / HC ratio at the input of step b) is thus the same for each reactor in parallel with step b). Examples
• Table 1 shows the characteristics of the feed (catalytic cracking gasolines) treated by the process according to the invention.
• the charge is heavy FCC gasoline.
• the analysis methods used to characterize the loads and effluents are as follows:
• Example 1 (comparative): hydrodesulphurization of gasoline on a catalyst allowing the desulphurization step a) according to the invention.
• the gasoline feed is treated by a desulfurization step a) according to the invention.
• Desulfurization step a) was carried out with 50 ml of CoMo / alumina catalyst, which are placed in an isothermal tubular reactor, with a fixed catalyst bed.
• the catalyst is first sulfurized by treatment for 4 hours under a pressure of 2 MPa at 350 ° C, in contact with a feed consisting of 2% by weight of sulfur in the form of dimethyldisulfide in n-heptane.
• Olefins (% weight) 20% 18%
• Table 2 Comparison of the characteristics of the feed and of the desulfurized gasoline according to step a) of the invention.
• ** % H OD denotes the rate of hydrogenation of olefins
• the desulfurized effluent contains more mercaptan type compounds than the feed because the mercaptans are produced by the recombination reactions between the olefins present in the feed and the hS produced by the hydrodesulphurization reactions.
• Example 2 (comparative): hydrodesulfurization of the total effluent obtained from Example 1 with a finishing hydrodesulfurization catalyst.
• the total effluent from step a) of desulfurization of Example 1 is subjected to finishing hydrodesulfurization.
• the total effluent from step a) consists of:
• the total effluent from step a) is subjected to finishing hydrodesulfurization on a nickel-based catalyst, in an isothermal tubular reactor, with a fixed catalyst bed.
• the finishing catalyst is prepared from a transition alumina of 140 m 2 / g in the form of beads 2 mm in diameter.
• the pore volume is 1 ml / g of support.
• 1 kilogram of support is impregnated with 1 liter of nickel nitrate solution.
• the catalyst is then dried at 120 ° C. and calcined in a stream of air at 400 ° C. for one hour.
• the nickel content of the catalyst is 20% by weight.
• the catalyst (100 ml) is then sulfurized by treatment for 4 hours under a pressure of 2 MPa at 350 ° C., in contact with a charge containing 2% by weight of sulfur in the form of dimethyldisulfide in n-heptane.
• the finishing hydrodesulphurization H2 / HC ratio is undergone because no addition of hydrogen is made between hydrodesulphurization step a) and the finishing hydrodesulphurization step.
• the test temperature is 380 ° C.
• the effluent is cooled and the condensed gasoline obtained after cooling is subjected to a hydrogen stripping step to rid the gasoline of dissolved H2S.
• the characteristics of the gasoline obtained after stripping are presented in Table 3.
• Olefins (% weight) 18% 18%
• Table 3 Characteristics of gasoline before and after finishing hydrodesulphurization on a nickel catalyst
• the gasoline treated with a finishing hydrodesulfurization of example 2 contains 7 ppm S in the form of mercaptans, which corresponds to a mercaptan desulfurization rate of 67%.
• the gasoline obtained has 14 ppm of total sulfur, which corresponds to a degree of desulfurization of the finishing step of 56%.
• the nickel-based catalyst makes it possible to desulphurize gasoline and reduce its mercaptan content without significantly hydrogenating the olefins in gasoline.
• the degree of hydrogenation of the olefins is negligible, this makes it possible to avoid a loss of octane in this stage.
• Example 3 (according to the invention): hydrodesulphurization of the total effluent from Example 1 with a finishing hydrodesulphurization catalyst and with the addition of hydrogen.
• the total effluent from step a) of desulfurization of Example 1 is subjected to a finishing hydrodesulfurization with an additional addition of hydrogen according to an embodiment of step b) of the invention.
• the total effluent from step a) consists of:
• the total effluent from step a) is subjected to finishing hydrodesulfurization with additional addition of hydrogen over a nickel-based catalyst.
• the nickel-based finishing catalyst is prepared in the same way as that used in the example
• Example 2 The catalyst is subjected to a sulfurization procedure identical to that described in Example 2.
• the addition of additional hydrogen to that which comes from the total effluent from step a) is then carried out so as to have an H2 / HC ratio at the inlet of the finishing hydrodesulfurization reactor of 697 expressed in liters of hydrogen at normal conditions per liter of feed at standard conditions.
• Olefins (% weight) 18% 18%
• This gasoline has 8 ppm S in the form of mercaptans, which corresponds to a mercaptan desulfurization rate of 62%.
• the nickel-based catalyst allows gasoline to be desulfurized and its mercaptan content to be reduced without significantly hydrogenating the olefins in gasoline.
• the degree of hydrogenation of the olefins is negligible, this makes it possible to avoid a loss of octane in this stage.
• the two gasolines obtained by a finishing hydrodesulphurization treatment (Example 2 and Example 3) have the same total sulfur content: 14 ppm by weight.
• the mercaptan content of these gasolines is also very similar (7 and 8 ppm S in the form of mercaptans respectively).
• the two gasolines therefore have very similar characteristics given that their contents of total sulfur, of sulfur in the form of mercaptans and also the content of olefins are all very similar.
• T 380 ° C
• F adjustment factor
• the increase in the H2 / HC ratio at the inlet of step b) according to the invention also has no effect on the loss of octane from the gasoline because the olefins at the inlet of the reactor b) are not hydrogenated with the nickel-based catalyst, even with an H2 / HC 1, 94 times higher than the base case. Consequently, the increase in the H2 / HC ratio at the input of step b) according to the invention does not generate degradation of the octane of the gasoline nor of overconsumption of hydrogen in the process.

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• 2019
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