Process for treating gasoline containing sulfur compounds comprising a dilution step

Abstract

A process for treating gasoline containing sulfur compounds, olefins and diolefins, the process comprising the following steps: a) bringing gasoline, hydrogen and a hydrodesulfurization catalyst into contact to obtain a partially desulfurized effluent; b) bringing the partially desulfurized effluent obtained at the end of step a) into contact directly and without any separation with a gaseous or liquid feedstock as a diluent, under normal temperature and pressure conditions, to obtain a partially desulfurized and diluted effluent; c) bringing the partially desulfurized and diluted effluent obtained at the end of step b) into contact with hydrogen and a catalyst to obtain a desulfurized effluent.

Description

Title of the invention: Method for treating gasoline containing sulfur compounds comprising a dilution step Field of the invention

[0001] The present invention relates to a process for producing gasoline with low sulfur and mercaptan content. State of the art

[0002] The production of gasolines meeting new environmental standards requires a significant reduction in their sulfur content.

[0003] It is also known that conversion gasolines, and more particularly those from catalytic cracking, which can represent 30 to 50% of the gasoline pool, have high mono-olefin and sulfur contents.

[0004] The sulfur present in gasolines is for this reason attributable, at nearly 90%, to gasolines resulting from catalytic cracking processes, which will be called in the following FCC gasolines (Fluid Catalytic Cracking according to the English terminology, which can be translated as fluidized bed catalytic cracking). FCC gasolines therefore constitute the preferred feedstock for the process of the present invention.

[0005] Among the possible ways to produce low-sulfur fuels, the one that has been very widely adopted consists of specifically treating sulfur-rich gasoline bases by catalytic hydrodesulfurization processes in the presence of hydrogen. Traditional processes desulfurize gasolines non-selectively by hydrogenating a large portion of the mono-olefins, which results in a high loss in octane number and high hydrogen consumption. The most recent processes, such as the Prime G+ process (trademark), make it possible to desulfurize olefin-rich cracked gasolines, while limiting the hydrogenation of the mono-olefins and consequently the loss of octane and the resulting high hydrogen consumption. Such processes are for example described in patent applications EP1077247 and EPI 174485.

[0006] Patent application US2009065396 discloses a multi-stage hydrodesulfurization process such that the second stage is operated at a higher temperature.

[0007] Documents EP1857527 and WO2001 / 38457 disclose two-stage hydrodesulfurization processes with an intermediate stage of extraction of the H2S formed during the first stage.

[0008] Documents US5985136 and WO2003 / 099963 disclose two-step processes in which the catalysts used in the first and second steps are different.

[0009] Nevertheless, there is still a need to limit losses by hydrogenation of olefins during hydrodesulfurization treatments of gasoline feedstocks to achieve sulfur specifications. Objects of the invention

[0010] The purpose of the pre-sale invention is to implement a process for producing low-sulfur gasolines, making it possible to upgrade the entirety of a gasoline cut containing sulfur, preferably a catalytic cracking gasoline cut, and to reduce the sulfur contents in said gasoline cut to very low levels, without any significant reduction in gasoline yield while minimizing the reduction in octane number due to the hydrogenation of olefins.

[0011] Thus, the subject of the present invention is a process for treating a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps:

[0012] a) 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 are brought into contact at a temperature of between 210 and 320°C, at a pressure of between 1 and 4 MPa, with a space velocity of between 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of feedstock to be treated expressed in m3 per hour under standard conditions of between 100 and 600 Nm3 / m3, to obtain a partially desulfurized effluent;

[0013] b) the partially desulfurized effluent obtained at the end of step a) is brought into contact directly and without any separation with a gaseous or liquid feedstock as diluent, under normal temperature and pressure conditions, defined here by a temperature of 15°C (288.15 K) and a pressure of 0.1 MPa, to obtain a partially desulfurized and diluted effluent;

[0014] c) the partially desulfurized and diluted effluent obtained at the end of step b) is brought into contact with 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 of between 210 and 320°C, at a pressure of between 1 and 4 MPa, with a space velocity of between 1 and 10 h 1 to obtain a desulfurized effluent.

[0015] The Applicant has surprisingly discovered that carrying out an intermediate step of dilution of a gasoline feedstock to be treated between two hydrodesulfurization steps makes it possible to significantly improve the stability of hydrodesulfurization performance by controlling the exothermicity of the hydrogenation of olefins and thus makes it possible to limit the phenomenon of deactivation of the hydrodesulfurization catalysts.

[0016] According to one or more embodiments, said diluent of step b) is a filler gaseous charge comprising at least 50% by volume of hydrogen relative to the total volume of said gaseous charge.

[0017] According to one or more embodiments, the ratio between the volume flow rate of said gaseous charge introduced in step b) and the volume flow rate of the hydrogen introduced in step a) is between 0.01 and 25 Nm3 / Nm3.

[0018] According to one or more embodiments, said diluent of step b) is a liquid hydrocarbon feedstock comprising a boiling temperature of between 30°C and 250°C.

[0019] According to one or more embodiments, the ratio between the volume flow rate of said liquid charge introduced in step b) and the volume flow rate of said gasoline charge introduced in step a) is between 0.01 and 1 SmVSm3 (Standard cubic meter per Standard cubic meter).

[0020] According to one or more embodiments, said liquid charge introduced in step b) is identical to the gasoline charge introduced in step a) of said method.

[0021] According to one or more embodiments, said gaseous feedstock or said liquid feedstock is introduced in step b) at a temperature lower than the temperature of the partially desulfurized effluent obtained at the end of step a).

[0022] According to one or more embodiments, said gaseous charge or said liquid charge is introduced in step b) at a temperature between 20°C and 300°C.

[0023] According to one or more embodiments, steps a) and b) are carried out in the same reactor.

[0024] According to one or more embodiments, steps a), b) and c) are carried out in the same reactor.

[0025] According to one or more embodiments, the catalyst of step a) and / or step c) comprises a group VIII metal content of between 0.1 and 10% by weight of group VIII metal oxide relative to the total weight of the catalyst, and a group VIB metal content of between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst.

[0026] According to one or more embodiments, the catalyst of step a) and / or step c) 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 catalyst of cobalt oxide, in CoO form, of between 0.1 and 10%, a content by weight relative to the total weight of catalyst of molybdenum oxide, in MoO3 form, of between 1 and 20%, a cobalt / molybdenum molar ratio of between 0.1 and 0.8, a content by weight relative to the total weight of catalyst of phosphorus oxide in P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface area of ​​between 50 and 250 m2 / g.

[0027] According to one or more embodiments, the catalysts of steps a) and c) are identical.

[0028] According to one or more embodiments, before step a) and before any possible distillation step, the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst to selectively hydrogenate the diolefins contained in said gasoline into olefins.

[0029] According to one or more embodiments, the gasoline is a catalytic cracking gasoline. List of figures

[0030] [Fig-1] [Fig.l] is a schematic representation of an embodiment according to the invention. Detailed Description Definitions

[0031] In the following, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUP AC classification.

[0032] Specific surface area means the BET specific surface area (SBet in m2 / g) determined by nitrogen adsorption in accordance with ASTM D 3663-78 established from the BRUNAUER-EMMETT-TELLER method described in the periodical "The Journal of American Society", 1938, 60, 309.

[0033] The total pore volume of the catalyst or support used for the preparation of the catalyst is understood to mean the volume measured by intrusion with a mercury porosimeter according to standard ASTM D4284 at a maximum pressure of 4000 bar (400 MPa), using a surface tension of 484 dyne / cm and a contact angle of 140°, for example with an Autopore III model device from the Microméritics® brand.

[0034] The wetting angle was taken equal to 140° following the recommendations of the work "Techniques de l'ingénieur, traité analyse et caractérisation", pages 1050-1055, written by Jean Charpin and Bernard Rasneur. In order to obtain better precision, the value of the total pore volume corresponds to the value of the total pore volume measured by intrusion with a mercury porosimeter measured on the sample minus the value of the total pore volume measured by intrusion with a mercury porosimeter measured on the same sample for a pressure corresponding to 30 psi (approximately 0.2 MPa).

[0035] The contents of elements of group VIII, group VIB and phosphorus are measured by X-ray fluorescence. Description of the load

[0036] 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 from a coking unit (coking according to English terminology), visbreaking (visbreaking according to English terminology), steam cracking (steam cracking according to English terminology) or catalytic cracking (FCC, Fluid Catalytic Cracking according to English terminology). This gasoline may optionally comprise a significant fraction of gasoline from other production processes such as atmospheric distillation (gasoline from direct distillation (or straight run gasoline according to English terminology)) or from conversion processes (coking or steam cracking gasoline). Said feedstock preferably consists of a gasoline cut from a catalytic cracking unit.

[0037] The feedstock is a gasoline cut containing sulfur compounds and olefins whose boiling point range 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 220°C. The process according to the invention can also treat feedstocks having end points lower than those mentioned above, such as for example a C5-180°C cut.

[0038] The sulfur content of gasoline cuts produced by catalytic cracking (FCC) depends on the sulfur content of the feedstock treated by the FCC, the presence or absence of pretreatment of the FCC feedstock, as well as the end point of the cut. Generally, the sulfur contents of the entire gasoline cut, in particular those from the FCC, are greater than 100 ppm by weight and most of the time greater than 500 ppm by weight. For gasolines having end points above 200°C, the sulfur contents are often greater than 1000 ppm by weight, and in some cases they can even reach values ​​of the order of 4000 to 5000 ppm by weight.

[0039] The feedstock treated by the process according to the invention may be a feedstock containing sulfur compounds in a content greater than 1000 ppm by weight of sulfur, and often greater than 1500 ppm.

[0040] Furthermore, gasolines from catalytic cracking units (FCC) contain, on average, between 0.5% and 5% by weight of diolefins, between 20% and 50% by weight of olefins, between 10 ppm and 0.5% by weight of sulfur, of which generally less than 300 ppm of mercaptans. Step aO) Selective hydrogenation (optional)

[0041] Depending on the type of gasoline to be treated, it may be advantageous to pre-treat the gasoline in the presence of hydrogen and a selective hydrogenation catalyst. so as to at least partially hydrogenate the diolefins and carry out a reaction to make part of the light mercaptan compounds (RSH) present in the thioether feed heavier, by reaction with olefins.

[0042] For this purpose, the gasoline to be treated is sent to a selective hydrogenation catalytic reactor containing at least one fixed or mobile bed of catalyst for selective hydrogenation of diolefins and for weighting of light mercaptans. The reaction for selective hydrogenation of diolefins and weighting of 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 VIB and an oxide support. The element from group VIII is preferably chosen from nickel and cobalt and in particular nickel. The element from group VIB, when present, is preferably chosen from molybdenum and tungsten and very preferably molybdenum.

[0043] The oxide support of the catalyst is preferably chosen from alumina, nickel aluminate, silica, silicon carbide, or a mixture of these oxides. Preferably, alumina is used, and even more preferably, high-purity alumina. According to a preferred embodiment, the selective hydrogenation catalyst contains nickel with a content by weight of nickel oxide, in NiO form, of between 1 and 12%, and molybdenum with a content by weight of molybdenum oxide, in MoO3 form, of between 6% 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 sulfurization rate of the metals constituting the catalyst is preferably greater than 60%.

[0044] During the optional selective hydrogenation step, the gasoline is brought into contact 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) between 0.5 h 1 and 20 h1, the unit of the liquid space velocity being the volume of charge per volume of catalytic bed and per hour (L / L / h). 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 m3 per hour and the flow rate of feedstock to be treated expressed in m3 per hour under standard conditions of between 2 and 100 Nm3 / m3, preferably between 3 and 30 Nm3 / m3.

[0045] After selective hydrogenation, the diolefin content, determined by means of the maleic anhydride index (MAV or "Maleic Anhydride Value" according to English terminology), according to the UOP 326 method, is generally reduced to less than 6 mg of maleic anhydride / g, or even less than 4 mg AM / g and more preferred less than 2 mg AM / g. In some cases, it can be obtained less than 1 mg AM / g.

[0046] The selectively hydrogenated gasoline can then be distilled into at least two cuts, a light cut and a heavy cut and optionally an intermediate cut. In the case of fractionation into two cuts, the heavy cut is treated according to the process of the invention. In the case of fractionation into three cuts, the intermediate and heavy cuts can be treated separately by the process according to the invention.

[0047] It should be noted that it is possible to carry out the steps of hydrogenation of the diolefins and fractionation into two or three cuts simultaneously by means of a catalytic distillation column which includes a distillation column equipped with at least one catalytic bed. Step a) Hydrodesulfurization

[0048] The hydrodesulfurization step a) is carried out to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H2S.

[0049] The temperature is generally between 210 and 320°C and preferably between 220 and 300°C. The temperature used must be sufficient to maintain the gasoline to be treated in the vapor phase in the reactor.

[0050] The operating pressure of this step is generally between 1 and 4 MPa and preferably between 1.5 and 3 MPa.

[0051] The quantity of catalyst used in each reactor is generally such that the ratio between the flow rate of gasoline to be treated expressed in m3 per hour at standard conditions, per m3 of catalytic bed (also called space velocity or LHSV) is between 1 and 100 h 1 and preferably between 1 and 50 h1, very preferably between 2 and 20 h1.

[0052] The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m3 per hour (Nm3 / h) and the flow rate of feedstock to be treated expressed in m3 per hour at standard conditions (15°C, 0.1 MPa) is between 10 and 1000 Nm3 / m3, preferably between 50 and 600 Nm3 / m3. Normal m3 is understood to mean the volume of 1 m3 of gas at 0°C and 0.1 MPa.

[0053] The hydrogen required for this step may be fresh hydrogen or recycled hydrogen, preferably freed from H2S, or a mixture of fresh hydrogen and recycled hydrogen. Preferably, a mixture of fresh hydrogen and recycled hydrogen will be used.

[0054] The desulfurization rate 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 from step a) contains less than 100 ppm by weight of sulfur and preferably less than 50 ppm by weight of sulfur.

[0055] The catalyst used in step a) must have good selectivity towards hydrodesulfurization reactions compared to the olefin hydrogenation reaction.

[0056] The hydrodesulfurization catalyst of step a) comprises an oxide support and an active phase comprising a group VIB metal and a group VIII metal and optionally phosphorus and / or an organic compound as described below.

[0057] The group VIB metal present in the active phase of the catalyst is preferably chosen from molybdenum and tungsten. The group VIII metal present in the active phase of the catalyst is preferably chosen from cobalt, nickel and the 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.

[0058] The content of group VIII metal is between 0.1 and 10% by weight of group VIII metal oxide relative to the total weight of the catalyst, preferably between 0.6 and 8% by weight, preferably between 0.6 and 7% by weight, very preferably between 1 and 6% by weight.

[0059] The content of group VIB metal is between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst, preferably between 2 and 18% by weight, very preferably between 3 and 16% by weight.

[0060] The molar ratio of group VIII metal to group VIB metal of the catalyst is generally between 0.1 and 0.8, preferably between 0.2 and 0.6.

[0061] Optionally, the catalyst may further have a phosphorus content generally between 0.3 and 10% by weight of P2O5 relative to the total weight of catalyst, preferably between 0.3 and 5% by weight, very preferably between 0.5 and 3% by weight. For example, the phosphorus present in the catalyst is combined with the metal from group VIB and optionally also with the metal from group VIII in the form of heteropolyanions.

[0062] Furthermore, 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.

[0063] Preferably, the catalyst is characterized by a specific surface area of ​​between 5 and 400 m2 / g, preferably between 10 and 250 m2 / g, preferably between 50 and 250 m2 / g. The specific surface area is determined in the present invention by the BET method according to standard ASTM D3663, as described in the work Rouquerol F.; Rouquerol J.; Singh K. “Adsorption by Powders & Porous Solids: Principle, methodology and applications”, Academie Press, 1999, for example using an Autopore III™ model device from the Microméritics™ brand.

[0064] The total pore volume of the catalyst is generally between 0.4 cmVg and 1.3 cmVg, preferably between 0.6 cmVg and 1.1 cmVg. The total pore volume is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140°, as described in the same work.

[0065] The packed filling density (DRT) of the catalyst is generally between 0.4 and 0.8 g / mL, preferably between 0.4 and 0.7 g / mL. The DRT measurement consists of introducing the catalyst into a test tube whose volume has been previously determined and then, by vibration, packing it until a constant volume is obtained. The apparent density of the packed product is calculated by comparing the mass introduced and the volume occupied after packing.

[0066] The catalyst can be in the form of small diameter extrudates, cylindrical or multi-lobed (tri-lobed, quadri-lobed, etc.), or spheres.

[0067] The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas or even titanium or magnesium oxides used alone or in a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, the family of transition aluminas and silica-aluminas, very preferably, the oxide support is essentially made 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 is preferably made solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, i.e. which contains aluminas of theta, delta, kappa or alpha phase, alone or in a mixture and a quantity of less than 20% by weight of alumina of gamma, chi or eta phase.

[0068] In the process according to the invention, the conversion of the unsaturated sulfur compounds is advantageously greater than 15% and preferably greater than 50%. At the same time, the hydrogenation rate of the olefins is preferably less than 50%, more preferably less than 40%, and very preferably less than 35%, during this step.

[0069] The partially desulfurized effluent obtained at the end of step a) is then sent directly and without separation to step b) of the process according to the invention. Step b) dilution of the effluent from step a)

[0070] This step consists of diluting the partially desulfurized effluent obtained at the end of step a) with a gaseous or liquid feedstock as diluent (also called here “dilution feedstock” or “diluent”) under normal temperature and pressure conditions, defined here by a temperature of 15°C (288.15 K) and a pressure of 0.1 MPa. Preferably, the dilution feedstock is gaseous under normal temperature and pressure conditions.

[0071] In one embodiment according to the invention, when the diluent is a filler gaseous, under normal temperature and pressure conditions, said gaseous charge preferably comprises more than 50% by volume of hydrogen H2, even more preferably more than 70% by volume of hydrogen, and very preferably more than 90% by volume of hydrogen. The volume flow rate of the gaseous charge is defined so that the ratio between the volume flow rate of said gaseous charge introduced in step b) and the volume flow rate of the hydrogen introduced in step a) is between 0.01 and 25 NmVNm3, preferably between 0.02 and 4 NmVNm3, and very preferably between 0.04 and 1 NmVNm3. The gaseous charge is introduced under a pressure generally between approximately 1 and approximately 4 MPa, preferably between 1.5 and 3 MPa.The gaseous feedstock is generally introduced at a temperature lower than the temperature of the effluent from step a), preferably at a temperature between 20°C and 300°C, preferably between 30°C and 280°C, more preferably between 30°C and 220°C, even more preferably between 35°C and 180°C, more preferably between 35°C and 120°C, and even more preferably between 35°C and 80°C.

[0072] In another embodiment according to the invention, when the diluent is a liquid feedstock, under normal temperature and pressure conditions defined here by a temperature of 15°C (288.15 K) and a pressure of 0.1 MPa, said liquid feedstock may be any hydrocarbon liquid feedstock comprising a boiling point of between 30°C and 250°C, preferably of between 35°C and 240°C, and even more preferably of between 40°C and 220°C. Preferably, the liquid feedstock is identical to the gasoline feedstock to be treated introduced during step a) or to the effluent from step c). The volume flow rate of said liquid charge is defined so that the ratio between the volume flow rate of said liquid charge introduced in step b) and the volume flow rate of the gasoline charge introduced in step a) is between 0.01 and 1 mVm 3, preferably between 0.02 and 1 m3 / m3, and very preferably between 0.02 and 0.5 m3 / m3.The liquid feedstock is introduced under a pressure generally between about 1 and about 4 MPa, preferably between 1.5 and 3 MPa. The liquid feedstock is introduced at a temperature lower than the temperature of the effluent from step a), and is preferably between 20°C and 300°C, preferably between 30°C and 280°C, very preferably between 40°C and 220°C, even more preferably between 40°C and 180°C, more preferably between 40°C and 120°C, and even more preferably between 40°C and 80°C.

[0073] Preferably, step b) is carried out in the same reactor as step a). Step c) Hydrodesulfurization

[0074] This step consists of transforming at least part of the sulfur compounds contained in the effluent from step b) such as thiophene compounds, into saturated compounds, for example thiophanes (or thiacyclopentanes) or mercaptans, or to at least partially hydrogenolyze these sulfur compounds to form H2S.

[0075] The temperature is generally between 210 and 320°C and preferably between 220 and 300°C. The temperature used must be sufficient to maintain the gasoline to be treated in the vapor phase in the reactor.

[0076] The operating pressure of this step is generally between 1 and 4 MPa and preferably between 1.5 and 3 MPa.

[0077] The quantity of catalyst used in each reactor is generally such that the ratio between the flow rate of gasoline to be treated expressed in m3 per hour at standard conditions, per m3 of catalytic bed (also called space velocity or LHSV) is between 1 and 100 h 1 and preferably between 1 and 50 h1, very preferably between 3 and 20 h *.

[0078] The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m3 per hour (Nm3 / h) and the flow rate of charge to be treated expressed in m3 per hour at standard conditions (15°C, 0.1 MPa) is between 50 and 1000 Nm3 / m3, preferably between 100 and 600 Nm3 / m.

[0079] The catalyst used in step c) must have good selectivity towards hydrodesulfurization reactions compared to the olefin hydrogenation reaction.

[0080] The hydrodesulfurization catalyst of step c) comprises an oxide support and an active phase comprising a group VIB metal and a group VIII metal and optionally phosphorus and / or an organic compound as described below.

[0081] The group VIB metal present in the active phase of the catalyst is preferably chosen from molybdenum and tungsten. The group VIII metal present in the active phase of the catalyst is preferably chosen from cobalt, nickel and the 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.

[0082] The content of group VIII metal is between 0.1 and 10% by weight of group VIII metal oxide relative to the total weight of the catalyst, preferably between 0.6 and 8% by weight, preferably between 0.6 and 7% by weight, very preferably between 1 and 6% by weight.

[0083] The content of group VIB metal is between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst, preferably between 2 and 18% by weight, very preferably between 3 and 16% by weight.

[0084] The molar ratio of group VIII metal to group VIB metal of the catalyst is generally between 0.1 and 0.8, preferably between 0.2 and 0.6.

[0085] Optionally, the catalyst may further have a phosphorus content gene- typically between 0.3 and 10% by weight of P2O5 relative to the total weight of catalyst, preferably between 0.3 and 5% by weight, very preferably between 0.5 and 3% by weight. For example, the phosphorus present in the catalyst is combined with the metal from group VIB and optionally also with the metal from group VIII in the form of heteropolyanions.

[0086] Furthermore, 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.

[0087] Preferably, the catalyst is characterized by a specific surface area of ​​between 5 and 400 m2 / g, preferably between 10 and 250 m2 / g, preferably between 50 and 250 m2 / g. The specific surface area is determined in the present invention by the BET method according to standard ASTM D3663, as described in the work Rouquerol F.; Rouquerol J.; Singh K. “Adsorption by Powders & Porous Solids: Principle, methodology and applications”, Academie Press, 1999, for example using an Autopore III™ model device from the Microméritics™ brand.

[0088] The total pore volume of the catalyst is generally between 0.4 cmVg and 1.3 cmVg, preferably between 0.6 cmVg and 1.1 cmVg. The total pore volume is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140°, as described in the same work.

[0089] The packed filling density (DRT) of the catalyst is generally between 0.4 and 0.8 g / mL, preferably between 0.4 and 0.7 g / mL. The DRT measurement consists of introducing the catalyst into a test tube whose volume has been previously determined and then, by vibration, packing it until a constant volume is obtained. The apparent density of the packed product is calculated by comparing the mass introduced and the volume occupied after packing.

[0090] The catalyst can be in the form of small diameter extrudates, cylindrical or multi-lobed (tri-lobed, quadri-lobed, etc.), or spheres.

[0091] The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas or even titanium or magnesium oxides used alone or in a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, the family of transition aluminas and silica-aluminas, very preferably, the oxide support is essentially made 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 is preferably made solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, i.e. which contains aluminas of theta, delta, kappa or alpha phase, alone or in a mixture and a quantity of less than 20% by weight of alumina of gamma, chi or eta phase.

[0092] In the process according to the invention, the conversion of the unsaturated sulfur compounds is advantageously greater than 15% and preferably greater than 50%. At the same time, the hydrogenation rate of the olefins is preferably less than 50%, more preferably less than 40%, and very preferably less than 35%, during this step.

[0093] In a preferred variant, the catalyst used during step c) is the same as that used during step a).

[0094] Preferably, step c) is carried out in the same reactor as step b).

[0095] Very preferably, steps a), b) and c) are carried out in the same reactor.

[0096] Step d) Finishing hydrodesulfurization (optional)

[0097] During hydrodesulfurization steps a) and c), a large part of the sulfur compounds is transformed into H2S. The remaining sulfur compounds are essentially refractory sulfur compounds and recombinant mercaptans resulting from the addition of the H2S formed in steps a) and c) to the monoolefins present in the feedstock.

[0098] The so-called finishing hydrodesulfurization step is mainly implemented to reduce the content of recombinant mercaptans. Preferably, step d) is carried out at a higher temperature than that of steps a) and c). Indeed, by using a higher temperature in this step compared to the temperature of steps a) and c), the formation of olefins and H2S will be favored by the thermodynamic equilibrium. Step d) also makes it possible to continue the hydrodesulfurization of the residual sulfur compounds.

[0099] The hydrodesulfurization step d) consists of bringing the effluent from step c) into contact, optionally with the addition of hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization.

[0100] The hydrodesulfurization step d) is carried out without significant hydrogenation of the olefins. The hydrogenation rate of the olefins of the catalyst of the hydrodesulfurization step d) is generally less than 5% and even more generally less than 2%.

[0101] The temperature of this step is generally between 280 and 400°C, more preferably between 290 and 380°C, and very preferably between 300 and 360°C. The temperature of this step d) is generally at least 5°C higher, preferably at least 10°C higher and very preferably at least 30°C higher than the temperature of steps a) and c).

[0102] The operating pressure of this step is generally between 0.5 and 5 MPa and preferably between 1 and 3 MPa.

[0103] The quantity of catalyst used in each reactor is generally such that the ratio between the flow rate of gasoline to be treated expressed in m3 per hour at standard conditions, per m3 of catalytic volume (also called space velocity or LHSV) is between 1 and 10 h 1 and preferably between 2 and 8 h1.

[0104] The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m3 per hour (Nm3 / h) and the flow rate of charge to be treated expressed in m3 per hour at standard conditions (15°C, 0.1 MPa) is between 50 and 600 Nm3 / m 3, preferably between 50 and 500 Nm3 / m3.

[0105] The desulfurization rate of step d), 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 from step d) contains less than 60 ppm by weight of sulfur and preferably less than 40 ppm by weight of sulfur.

[0106] The catalyst of step d) is of a different nature and / or composition from that used in steps a) and c). The catalyst of step d) is in particular a very selective hydrodesulfurization catalyst: it makes it possible to hydrodesulfurize without hydrogenating the olefins and therefore to maintain the octane number.

[0107] The catalyst which may be suitable for this step d) of the process according to the invention, without this list being exhaustive, is a catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, and preferably chosen from the group formed by nickel, cobalt, iron. These metals can be used alone or in combination. Preferably, the active phase consists of a metal from group VIII, preferably nickel. Particularly preferably, the active phase consists of nickel.

[0108] The content of group VIII metal is between 1 and 60% by weight of group VIII metal oxide relative to the total weight of the catalyst, preferably between 5 and 30% by weight, very preferably between 5 and 20% by weight.

[0109] Preferably, the catalyst is characterized by a specific surface area of ​​between 5 and 400 m2 / g, preferably between 10 and 250 m2 / g, preferably between 20 and 200 m2 / g, very preferably between 30 and 180 m2 / g. The specific surface area 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 Solids: Principle, methodology and applications”, Academie Press, 1999, for example using an Autopore III™ model device from the Microméritics™ brand.

[0110] The pore volume of the catalyst is generally between 0.4 cm3 / g and 1.3 cm3 / g, preferably between 0.6 cm3 / g and 1.1 cm3 / g. The total pore volume is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140°, as described in the same work.

[0111] The packed filling density (DRT) of the catalyst is generally understood between 0.4 and 0.8 g / mL, preferably between 0.4 and 0.7 g / mL.

[0112] The DRT measurement consists of introducing the catalyst into a test tube whose volume has been previously determined and then, by vibration, tamping it until a constant volume is obtained. The apparent density of the tamped product is calculated by comparing the mass introduced and the volume occupied after tamping.

[0113] The catalyst can be in the form of small diameter extrudates, cylindrical or multi-lobed (tri-lobed, quadri-lobed, etc.), or spheres.

[0114] The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas or even titanium or magnesium oxides used alone or in a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, the family of transition aluminas and silica-aluminas, very preferably, the oxide support is essentially made 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 is preferably made solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, i.e. which contains aluminas of theta, delta, kappa or alpha phase, alone or in a mixture and a quantity of less than 20% of alumina of gamma, chi or eta phase.

[0115] A very preferred embodiment of the invention corresponds to the use for step d) 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 NiO form, of between 5 and 20%, said catalyst having a specific surface area of ​​between 30 and 180 m2 / g.

[0116] The catalyst of the hydrodesulfurization step d) is characterized by a catalytic hydrodesulfurization 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 the hydrodesulfurization steps a) and c).

[0117] Description of the preparation of catalysts and sulfurization

[0118] The preparation of the catalysts of steps a), c) and d) is known and generally comprises a step of impregnation of the metals of group VIII and group VIB when present, and optionally of the phosphorus and / or of the organic compound on the oxide support, followed by drying, then by an optional calcination making it possible to obtain the active phase in their oxide forms. Before its use in a process for hydrodesulfurization of an olefinic gasoline cut containing sulfur, the catalysts are generally subjected to sulfurization in order to form the active species as described below.

[0119] The impregnation step can be carried out either by slurry impregnation or by excess impregnation, either 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 solubilize the metal precursors in the desired concentrations.

[0120] For example, among the sources of molybdenum, it is possible to use oxides and hydroxides, molybdic acids and their salts, in particular ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PMoi204o), and their salts, and possibly silicomolybdic acid (H4 SiMoi204o) and its salts. The sources of molybdenum can also be any heteropolycompound of the Keggin, lacunar Keggin, substituted Keggin, Dawson, Anderson, Strandberg type, for example. Molybdenum trioxide and heteropolycompounds of the Keggin, lacunar Keggin, substituted Keggin and Strandberg type are preferably used.

[0121] The tungsten precursors that can be used are also well known to those skilled in the art. For example, among the sources of tungsten, it is possible to use oxides and hydroxides, tungstic acids and their salts, in particular ammonium salts such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and their salts, and possibly silicotungstic acid (H4SiWi2O40) and its salts. The sources of tungsten can also be any heteropolycompound of the Keggin, vacated Keggin, substituted Keggin, Dawson type, for example. Preferably, ammonium oxides and salts such as ammonium metatungstate or heteropolyanions of the Keggin, vacated Keggin or substituted Keggin type are used.

[0122] The cobalt precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example. Cobalt hydroxide and cobalt carbonate are preferably used.

[0123] The nickel precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example.

[0124] The preferred phosphorus precursor is orthophosphoric acid H3PO4, but its salts and esters such as ammonium phosphates are also suitable. Phosphorus may also be introduced together with the group VIB element(s) in the form of Keggin, vacated Keggin, substituted Keggin or Strandberg type heteropolyanions.

[0125] After the impregnation step, 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, 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 reduced pressure. Preferably, this step is carried out at atmospheric pressure. It is advantageously carried out in a flow-through bed using air or any other hot gas. Preferably, when drying is carried out in a fixed bed, the gas used is either air or an inert gas such as argon or nitrogen. Very preferably, drying is carried out in a flow-through bed in the presence of nitrogen and / or air. Preferably, the drying step has a duration of between 5 minutes and 15 hours, preferably between 30 minutes and 12 hours.

[0126] According to a variant of the invention, 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 temperature above 200°C under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.

[0127] According to another preferred variant of the invention, 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 of between 250 and 1000°C and preferably between 200 and 750°C, for a duration typically of between 15 minutes and 10 hours, under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.

[0128] Before contacting with the feedstock to be treated in a gasoline hydrodesulfurization process, the catalysts of the process according to the invention generally undergo a sulfurization step. The sulfurization is preferably carried out in a sulforeducing medium, that is to say in the presence of H2S and hydrogen, in order to transform the metal oxides into sulfides such as, for example, MoS2, Co9S8 or Ni3S2. The sulfurization is carried out by injecting onto the catalyst a stream containing H2S and hydrogen, or a sulfur compound capable of decomposing into H2S in the presence of the catalyst and hydrogen. Polysulfides such as dimethyl disulfide (DMDS) are H2S precursors commonly used to sulfurize catalysts. The sulfur can also come from the feedstock. The temperature is adjusted so that the H2S reacts with the metal oxides to form metal sulfides.This sulfurization can be carried out in situ or ex situ (inside or outside the reactor) of the reactor of the process according to the invention at temperatures between 200 and 600°C, and more preferably between 300 and 500°C.

[0129] The sulfurization rate 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 sulfurized when the overall sulfurization 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 sulfurization of the metal(s) considered.

[0130]

[0131]

[0132]

[0133]

[0134]

[0135]

[0136]

[0137]

[0138] The overall sulfurization rate is defined by the following equation: (S / metal)cataiy seur > 0.6 x (S / metal) theoretical in which: (S / metal)catalyst is the molar ratio between sulfur (S) and metal present on the catalyst (S / metal) theoretical is the molar ratio between sulfur and metal corresponding to the total sulfurization of the metal into sulfide. This theoretical molar ratio varies depending on the metal considered: (S / Fe)theoretical 1 (S / Co)theoretical 1 (S / Ni)theoretical 1 - (S / Mo)theoretical=2 / 1 - (S / W)theoretical=2 / 1 When the catalyst comprises several metals, 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 sulfurization of each metal into sulfide, the calculation being carried out in proportion to the relative molar fractions of each metal. Other characteristics and advantages of the invention will now appear on reading the description which follows, given solely for illustrative and non-limiting purposes, and with reference to [Fig.l]. With reference to [Fig.l] and according to one embodiment of the process according to the invention, the gasoline to be treated is sent via line 1 and hydrogen via line 2 to the hydrodesulfurization unit 3 of step a). The hydrodesulfurization unit 3 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 fixed bed reactor is used. 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). During the hydrodesulfurization in step a), recombinant mercaptans are formed by addition of H2S formed on the olefins.The effluent from hydrodesulfurization unit 3 is then mixed with a diluent via line 5 in step b) and introduced into hydrodesulfurization unit 6 via line 4 without removing the H2S formed. The hydrodesulfurization unit 6 of step c) 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 fixed bed reactor is used. 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). During hydrodesulfurization in step c), recombinant mercaptans are formed by addition of H2S formed on the olefins. The effluent from hydrodesulfurization unit 6 is then introduced into the so-called finishing hydrodesulfurization unit 8 via line 7 without removing the H2S formed. The hydrodesulfurization unit 8 of step d) is for example a reactor containing a fixed-bed or fluidized-bed hydrodesulfurization catalyst; preferably, a fixed-bed reactor is used. Unit 8 is operated at a higher temperature than unit 6 and in the presence of a selective catalyst comprising an oxide support and an active phase consisting of at least one group VIII metal to decompose at least partially the recombinant mercaptans into olefins and H2S. It also allows the hydrodesulfurization of more refractory sulfur compounds.

[0139] An effluent is withdrawn from said hydrodesulfurization reactor 8 via line 9, a gasoline containing H2S, compounds having a boiling point higher than that of butane, and sulfur compounds including mercaptans.

[0140] The following examples illustrate the invention without limiting its scope. Examples

[0141] The analysis methods used to characterize the loads and effluents are as follows:

[0142] - density according to the NF EN ISO 12185 method;

[0143] - sulfur content according to ASTM D2622 method for contents above 10 ppm S and ISO 20846 for contents below 10 ppm S;

[0144] - distillation according to the CSD method simulated distillation "CSD according to the method ASTM2887;

[0145] - diolefin content, determined via the anhydride index maleic (MAV or “Maleic Anhydride Value” according to Anglo-Saxon terminology), according to the UOP 326 method. Example 1 (non-compliant)

[0146] A gasoline from a catalytic cracking unit composed of 25% by weight of olefins and 600 ppmS of total sulfur is subjected to a single hydrodesulfurization step in a single catalytic bed in an adiabatic reactor in the presence of a CoMo type catalyst A supported on alumina, the metal contents being respectively 3% by weight of CoO and 10% by weight of MoO3, the specific surface area of ​​the catalyst is 135 m2 / g. Prior to its use, the catalyst is sulfurized by treatment for 4 hours under a pressure of 3.4 MPa at 350°C, in contact with a charge consisting of 2% by weight of sulfur in the form of dimethyl disulfide in n-heptane.

[0147] 60 mL of the catalyst are loaded into the reactor in a single catalytic bed.

[0148] The operating conditions of the single-stage hydrodesulfurization stage of the gasoline feedstock are as follows: LHSV = 3 h 1 relative to the entire catalytic bed, H2 / HC = 250 NmVm3, P = 2.0 MPa.

[0149] The characteristics of the liquid effluent obtained at the end of the hydrodesulfurization step and the temperature conditions as a function of time under the load flow thus obtained are presented in Table 1.

[0150] [Tables 1] Time under load flow days 0 15 30 60 90 Inlet sulfur content ppmS 598 598 598 598 598 Outlet sulfur content ppmS 10 11 11 13 15 Inlet temperature °C 266 266 266 266 266 Outlet temperature °C 288 286 285 284 283

[0151] * Hydrogenation of olefins Example 2 (compliant)

[0152] Gasoline from a catalytic cracking unit whose characteristics are described in Example 1 is subjected to hydrodesulfurization on the sulfurized catalyst A described in Example 1. 60 mL of the catalyst are loaded into the reactor in two catalytic beds of 30 mL each.

[0153] The overall operating conditions of the hydrodesulfurization with respect to the set of two catalytic beds are as follows: LHSV = 3 h1, H2 / HC = 250 NmVm3, P = 2.0 MPa.

[0154] In the first catalytic bed, the operating conditions of the first hydrodesulfurization stage are as follows: LHSV = 6 h1, H2 / HC = 225 NmVm3.

[0155] An addition of gaseous diluent composed of hydrogen at the end of this step is carried out with a diluent flow rate / H2 flow rate ratio at the inlet of the 1st reactor = 0.22 NmVNm3. The diluent is added at a temperature of 50°C.

[0156] This mixture is then subjected to a second hydrodesulfurization step on catalyst A in the second catalyst bed. The operating conditions of this hydrodesulfurization step in the second bed are as follows: LHSV = 6 h1, H2 / HC = 275 NmVm3.

[0157] The characteristics of the liquid effluent from the second hydrodesulfurization stage and the temperature conditions as a function of time under the load flow thus obtained are presented in Table 2.

[0158] [Tables2] Time under load flow days 0 15 30 60 90 Inlet sulfur content ppmS 598 598 598 598 598 Outlet sulfur content ppmS 10 10 11 12 13 Bed 1 inlet temperature °C 269 269 269 269 269 Bed 1 outlet temperature °C 284 284 283 282 281 Bed 2 inlet temperature °C 275 275 274 273 272 Bed 2 outlet temperature °C 284 284 283 281 279

[0159] * Hydrogenation of olefins

[0160] The exotherm is lower in this example compared to Example 1 as well as the maximum temperature. This leads to slower deactivation of the catalyst and reduced sulfur contents at the outlet of the process according to the invention. Example 3 (compliant)

[0161] Gasoline from a catalytic cracking unit whose characteristics are described in Example 1 is subjected to hydrodesulfurization on the sulfurized catalyst A described in Example 1. 60 mL of the catalyst are loaded into the reactor in two beds of 30 mL.

[0162] The overall operating conditions of the hydrodesulfurization with respect to all the catalytic beds are as follows: LHSV = 3 h1, H2 / HC = 250 Nm3 / m3, P = 2.0 MPa.

[0163] In the first bed, the operating conditions of the hydrodesulfurization stage are as follows: LHSV = 6 h1, H2 / HC = 237.5 NmVm3.

[0164] An addition of gaseous diluent composed of hydrogen at the end of this step with a diluent flow rate / H2 flow rate ratio at the inlet of the 1st reactor = 0.105 Nm3 / Nm3. The diluent is added at a temperature of 50°C.

[0165] This mixture is then subjected to a second hydrodesulfurization step on catalyst A in the second catalyst bed. The operating conditions of this hydrodesulfurization step in the second bed are as follows: LHSV = 6 h1, H2 / HC = 262.5 Nm3 / m3.

[0166] The characteristics of the liquid effluent from the second hydrodesulfurization stage and the temperature conditions as a function of time under the load flow thus obtained are presented in Table 3.

[0167] [Tables3] Time under load flow days 0 15 30 60 90 Inlet sulfur content ppmS 598 598 598 598 598 Outlet sulfur content ppmS 10 10 11 12 13 Bed 1 inlet temperature °C 267 267 267 267 267 Bed 1 outlet temperature °C 281 280 280 279 278 Bed 2 inlet temperature °C 276 276 275 274 273 Bed 2 outlet temperature °C 286 285 285 283 281

[0168] * Hydrogenation of olefins

[0169] The exotherm is lower in this example compared to Example 1 as well as the maximum temperature. This leads to slower deactivation of the catalyst and reduced sulfur contents at the outlet of the process according to the invention.

Claims

Claims

1. A method for treating a gasoline containing sulfur compounds, olefins and diolefins, the method comprising at least the following steps: a) bringing the 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 into contact, at a temperature of between 210 and 320°C, at a pressure of between 1 and 4 MPa, with a space velocity of between 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of feedstock to be treated expressed in m3 per hour at standard conditions of between 100 and 600 Nm3 / m3, to obtain a partially desulfurized effluent;b) the partially desulfurized effluent obtained at the end of step a) is brought into contact directly and without any separation with a gaseous or liquid feedstock as diluent, under normal temperature and pressure conditions, the ratio between the volume flow rate of said gaseous feedstock introduced in step b) and the volume flow rate of the hydrogen introduced in step a) is between 0.04 and 1 Nm3 / Nm3, to obtain a partially desulfurized and diluted effluent; c) the partially desulfurized and diluted effluent obtained at the end of step b) is brought into contact with 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 h 1 to obtain a desulfurized effluent.;

2. The method of claim 1, wherein said diluent of step b) is a gaseous feed comprising at least 50% by volume of hydrogen relative to the total volume of said gaseous feed.

3. A process according to claim 1, wherein said diluent of step b) is a liquid hydrocarbon feedstock having a boiling point between 30°C and 250°C.

4. Method according to any one of claims 1 and 3, in which the ratio between the volume flow rate of said liquid charge introduced in step b) and the volume flow rate of said gasoline charge introduced in step a) is between 0.01 and 1 Sm3 / Sm3.

5. Method according to one of claims 3 or 4, in which said liquid charge introduced in step b) is identical to the gasoline charge introduced in step a) of said method.

6. Process according to any one of claims 1 to 5, wherein in step b) said gaseous feedstock or said liquid feedstock is introduced at a temperature lower than the temperature of the partially desulfurized effluent obtained at the end of step a).

7. A method according to claim 6, wherein said gaseous feedstock or said liquid feedstock is introduced in step b) at a temperature between 20°C and 300°C.

8. Process according to any one of claims 1 to 7, in which steps a) and b) are carried out in the same reactor.

9. Process according to any one of claims 1 to 8, in which steps a), b) and c) are carried out in the same reactor.

10. A process according to any one of claims 1 to 9, wherein the catalyst of step a) and / or step c) comprises a group VIII metal content of between 0.1 and 10% by weight of group VIII metal oxide relative to the total weight of the catalyst, and a group VIB metal content of between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst.

11. A process according to any one of claims 1 to 10, wherein the catalyst of step a) and / or step c) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight relative to the total catalyst weight of cobalt oxide, in CoO form, of between 0.1 and 10%, a content by weight relative to the total catalyst weight of molybdenum oxide, in MoO3 form, of between 1 and 20%, a cobalt / molybdenum molar ratio of between 0.1 and 0.8, a content by weight relative to the total catalyst weight of phosphorus oxide in P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface area of ​​between 50 and 250 m2 / g.

12. A process according to any one of claims 1 to 11, wherein the catalysts of steps a) and c) are identical.

13. A process according to any one of claims 1 to 12, wherein before step a) and before any possible distillation step, the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst to selectively hydrogenate the diolefins contained in said gasoline into olefins.

14. A process according to any one of claims 1 to 13, wherein the gasoline is a catalytic cracked gasoline.