Process for refinement hydrodesulfurization of gasoline using catalyst based on group VIB and group VIII metals and phosphorus on alumina support having low specific surface area

Abstract

Process for the treatment of gasoline containing sulfur compounds and olefins, said process comprising at least the steps of: a) contacting gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a Group VIB metal and a Group VIII metal, at least in part in sulfide form, and an oxide support, to obtain a partially desulfurized effluent; b) bringing the partially desulfurized effluent obtained at the end of step a) into direct contact, without separating the H2S formed in step a), with a catalyst comprising an active phase comprising at least one Group VIB metal and at least one Group VIII metal, at least in part in the form of sulfides, phosphorus, and an alumina-based porous support, the catalyst has a specific surface area greater than or equal to 20 m2 / g and less than 150 m2 / g.

Description

Field of the Invention

[0001] The present invention relates to the field of hydrotreating gasoline fractions, in particular gasoline fractions produced by fluidized bed catalytic cracking units. More particularly, the present invention relates to the use of catalysts in a process for preparing low-sulfur gasoline. The present invention is very particularly applicable to treating gasoline fractions containing olefins and sulfur, such as gasoline produced by catalytic cracking, for which it is desired to reduce the content of sulfur compounds without hydrogenating the olefins and aromatics. Background of the Invention

[0002] Automotive fuel specifications require a significant reduction in the sulfur content of these fuels, particularly gasoline. This reduction particularly aims to limit the content of sulfur and nitrogen oxides in motor vehicle exhaust gases. The specifications for gasoline fuels currently in force in Europe since 2009 provide for a maximum sulfur content of 10 weight ppm (parts per million). Such specifications are also in force in other countries such as the United States and China, which have required the same maximum sulfur content since January 2017. To meet these specifications, gasoline must be treated via a desulfurization process.

[0003] The main source of sulfur in gasoline bases is "cracked" gasoline, and in particular the gasoline fraction of the catalytic cracking process of vacuum distillates obtained from crude oil or atmospheric or vacuum distillation residues. The gasoline fraction from catalytic cracking, which on average represents 40% of the gasoline base, actually accounts for more than 90% of the sulfur in gasoline. Therefore, the preparation of low-sulfur gasoline requires a desulfurization step for catalytic cracking gasoline. Among the other sources of gasoline that may contain sulfur, mention may also be made of coker gasoline, visbreaking gasoline, or to a lesser extent, gasoline obtained from atmospheric distillation or steam cracking gasoline.

[0004] Removing sulfur from gasoline fractions involves specifically treating these sulfur-rich gasolines via a desulfurization process in the presence of hydrogen. This is thus called a hydrodesulfurization (HDS) process. However, these gasoline fractions, and more particularly fluid catalytic cracking (FCC) gasoline, contain a large proportion of unsaturated compounds in the form of monoolefins (from approximately 20% to 50% by weight), diolefins (from 0.5% to 5% by weight) and aromatics that contribute to obtaining a good octane number. These unsaturated compounds are unstable and react during the hydrodesulfurization treatment. Diolefins polymerize to form gums during the hydrodesulfurization treatment. The formation of such gums leads to the progressive deactivation of the hydrodesulfurization catalyst or the progressive clogging of the reactor. Therefore, before any treatment of these gasolines, the diolefins must be removed by hydrogenation. Conventional treatment processes desulfurize gasoline non-selectively by hydrogenating most of the monoolefins, which results in a high loss of octane number and high hydrogen consumption. The latest hydrodesulfurization processes make it possible to desulfurize cracked gasoline rich in monoolefins while limiting the hydrogenation of monoolefins and thus limiting the loss of octane. Such processes are described, for example, in the documents EP-A-1077247 and EP-A-1174485.

[0005] However, when very thorough desulfurization of cracked gasoline is required, a part of the olefins present in the cracked gasoline are hydrogenated on the one hand, and on the other hand, they recombine with H2S to form mercaptans. Such compounds with the chemical formula R-SH (where R is an alkyl group) are generally referred to as recombined mercaptans and usually account for 20 wt% to 80 wt% of the residual sulfur in the desulfurized gasoline. The reduction of the recombined mercaptan content can be achieved by catalytic hydrodesulfurization, but this leads to the hydrogenation of most of the monoolefins present in the gasoline, which in turn results in a significant reduction in the octane number of the gasoline and excessive consumption of hydrogen. In addition, it is known that the lower the target sulfur content, that is, when attempting to thoroughly remove the sulfur compounds present in the feedstock, the greater the proportionate octane loss caused by the hydrogenation of monoolefins during the hydrodesulfurization step.

[0006] Thus, gasoline can be treated by a sequence of two reactors as described in document EP 1 077 247; the purpose of the first stage (also called the selective HDS stage) is generally to carry out deep desulfurization of the gasoline with minimal olefin saturation (and no aromatic loss), thereby achieving the maximum octane retention rate. The catalyst used is generally a CoMo-type catalyst. During this stage, new sulfur compounds: recombined mercaptans are formed by the recombination of H2S from the desulfurization with olefins.

[0007] The second stage generally serves to minimize the amount of recombined mercaptans. In the second stage, the temperature is generally higher in order to thermodynamically promote the removal of the mercaptans. In practice, a furnace is thus placed between the two reactors in order to be able to raise the temperature of the second reactor above that of the first reactor.

[0008] The catalyst used in the refining process must be particularly selective so as not to cause olefin saturation (and no aromatic loss), thereby resulting in octane loss. Therefore, it must enable the reduction of the total sulfur and mercaptan content in the hydrocarbon fraction, preferably the gasoline fraction, to extremely low levels while minimizing the reduction of the octane number. Generally, the catalyst used is based on nickel.

[0009] It is known from patent FR 3,023,184 to provide a hydrotreating catalyst on an alumina-based support, said hydrotreating catalyst comprising at least a Group VIB metal, a Group VII metal and phosphorus, and having a specific surface area of 20 to 150 m 2 / g.

[0010] It is also known from patent FR 2,840,315 to provide the use of a catalyst comprising at least one Group VIB metal, a Group VII metal and a support having a specific surface area of less than 200 m 2 / g.

[0011] However, it is still necessary to maximize the performance of the hydrotreated gasoline fraction to meet the sulfur specifications.

[0012] Surprisingly, the applicant has determined that when using a catalyst based on at least one Group VIII metal, at least one Group VIB metal, and phosphorus on an alumina support (which has a low specific surface area within a defined range) in a finishing hydrodesulfurization section (FNS) downstream of a selective hydrodesulfurization section (HDS), it is possible to reduce the processing temperature in gasoline hydrotreating while maintaining the process desulfurization performance and olefin retention (to maintain the octane number). Reducing the average processing temperature of the HDS section enables improvement of the overall cycle time of the process.

[0013] Without being bound by any theory, using a specific type of catalyst in finishing hydrodesulfurization enables removal of a portion of the refractory sulfur compounds in the finishing section while retaining olefins, which is brought about by the high selectivity achieved by the specific interaction between the active phase and the surface of the alumina support with a low specific surface area, thereby better controlling the olefin hydrogenation reaction and the conversion of sulfur compounds in the selective hydrodesulfurization and finishing hydrodesulfurization sections while maintaining as low a processing temperature as possible. Summary of the Invention

[0014] The object of the present invention is to implement a method for preparing gasoline with a low sulfur content, which can upgrade the entire sulfur-containing gasoline fraction, preferably the catalytic cracking gasoline fraction, and reduce the sulfur content in the gasoline fraction to a very low level without reducing the gasoline yield, while minimizing the reduction of the octane number caused by olefin hydrogenation.

[0015] The subject of the present invention is a method for treating gasoline containing sulfur compounds and olefins, the method comprising at least the following steps:

[0016] a) contacting gasoline, hydrogen, and a hydrodesulfurization catalyst at a temperature of 200 °C to 350 °C and a pressure of 0.2 MPa to 5 MPa, with a space velocity of 1 h -1 to 20 h -1 , where the ratio of the hydrogen flow rate expressed as standard m 3 / h to the flow rate of the feedstock to be treated expressed as m 3 / h under standard conditions is 10 Nm 3 / m 3 to 1000 Nm 3 / m 3 , and the hydrodesulfurization catalyst comprises an active phase and an oxide support, and the active phase comprises a Group VIB metal and a Group VIII metal at least partially in the sulfide form, thereby obtaining a partially desulfurized effluent;

[0017] b) Without separating the H2S formed in step a), directly contacting the partially desulfurized effluent obtained at the end of step a) with a refined hydrodesulfurization catalyst at a temperature of 250 °C to 400 °C and a pressure of 0.2 MPa to 5 MPa, where the space velocity is 1 h -1 to 20 h -1 , thereby obtaining a desulfurized effluent. The refined hydrodesulfurization catalyst comprises an active phase and a porous alumina-based support. The active phase comprises at least one Group VIB metal and at least one Group VIII metal at least partially in the form of sulfides, and phosphorus, preferably consisting of at least one Group VIB metal and at least one Group VIII metal at least partially in the form of sulfides and phosphorus. The content of the Group VIB metal in oxide form is 1 wt% to 20 wt% relative to the total weight of the catalyst, the content of the Group VIII metal in oxide form is 0.2 wt% to 10 wt% relative to the total weight of the catalyst, and the content of phosphorus in the form of P2O5 is 0.1 wt% to 5 wt% relative to the total weight of the catalyst. The catalyst has a specific surface area greater than or equal to 20 m 2 / g and less than 150 m 2 / g.

[0018] According to one or more embodiments, the molar ratio of the Group VIII metal to the Group VIB metal in the active phase of the catalyst in step b) is 0.1 to 2.0 mol / mol.

[0019] According to one or more embodiments, the molar ratio of phosphorus to the Group VIB metal in the active phase of the catalyst in step b) is 0.1 to 2.0 mol / mol.

[0020] According to one or more embodiments, the specific surface area of the catalyst in step b) is 20 m 2 / g to 110 m 2 / g.

[0021] According to one or more embodiments, the catalyst in step b) comprises cobalt as the Group VIII metal and molybdenum as the Group VIB metal.

[0022] According to one or more embodiments, the catalyst of step b) comprises an active phase composed of molybdenum, cobalt, and phosphorus, and a porous alumina-based support. The cobalt content in the form of CoO oxide is 0.5 wt% to 5 wt% relative to the total weight of the catalyst, the molybdenum content in the form of MoO3 oxide is 3 wt% to 12 wt% relative to the total weight of the catalyst, the phosphorus content in the form of P2O5 oxide is 0.3 wt% to 3 wt% relative to the total weight of the catalyst, the molar ratio of cobalt to molybdenum is 0.3 to 1.0 mol / mol, the molar ratio of phosphorus to molybdenum is 0.2 to 0.5 mol / mol, and the specific surface area of the catalyst is 25 to 90 m 2 / g.

[0023] According to one or more embodiments, the catalyst of step a) comprises an alumina support and an active phase containing cobalt and molybdenum. The catalyst contains a weight content of cobalt oxide in the form of CoO of 0.1 wt% to 10 wt% relative to the total weight of the catalyst, and a weight content of molybdenum oxide in the form of MoO3 of 1 wt% to 20 wt% relative to the total weight of the catalyst, wherein the cobalt / molybdenum molar ratio is 0.1 to 0.8 mol / mol.

[0024] According to one or more embodiments, the catalyst of step a) has a specific surface area of 60 to 250 m 2 / g.

[0025] According to one or more embodiments, the temperature of step b) is higher than the temperature of step a).

[0026] According to one or more embodiments, the gasoline is catalytic cracking gasoline. Detailed Description

[0027] Definition

[0028] Hereinafter, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, edited by D.R. Lide, 81st edition, 2000 - 2001). For example, Group VIII (or Group VIIIB) according to the CAS classification corresponds to the metals in columns 8, 9, and 10 according to the new IUPAC classification.

[0029] The BET specific surface area is measured by nitrogen physical adsorption according to the standard ASTM D3663-03, and the method is described in the work "Adsorption by Powders & Porous Solids: Principles, Methodology and Applications" by Rouquerol F., Rouquerol J. and Singh K., Academic Press, 1999.

[0030] In the following description of the present invention, the "total pore volume" of the oxide support or catalyst is understood to refer to the volume measured by mercury porosimetry according to the standard ASTM D4284-83 at a maximum pressure of 4000 bar (400 MPa), using a surface tension of 484 dynes / cm and a contact angle of 140°. According to the recommendation on page 1050-5 of the publication "Techniques de l'ingénieur, traité analyse et caractérisation" [Techniques of the Engineer, Analysis and Characterization Treatise] written by Jean Charpin and Bernard Rasneur, the wetting angle is taken to be 140°.

[0031] For better accuracy, the values of the total pore volume given hereinafter in ml / g or in cm 3 / g correspond to the total mercury volume value measured in ml / g or in cm 3 / g on the said sample minus the mercury volume value measured in ml / g or in cm 3 / g on the same sample corresponding to a pressure of 30 psi (about 0.2 MPa).

[0032] The contents of Group VIII metals, Group VIB metals and phosphorus are measured by X-ray fluorescence.

[0033] The contents of Group VIB metals, Group VIII metals and phosphorus in the said catalyst are expressed as oxides after correcting for the loss on ignition of the catalyst sample in a muffle furnace at 550 °C for 2 h. The loss on ignition is due to the loss of moisture. It is determined according to ASTM D7348.

[0034] Raw materials

[0035] The method according to the invention enables any type of gasoline fraction containing sulfur compounds and olefins, either alone or as a mixture, such as fractions from coking, visbreaking, steam cracking or fluid catalytic cracking (FCC) units. This gasoline can optionally consist of a large proportion of gasoline originating from other production processes, such as atmospheric distillation (gasoline from direct distillation (or straight-run gasoline)) or conversion processes (coker or steam cracking gasoline). The feedstock preferably consists of a gasoline fraction produced by a fluid catalytic cracking unit.

[0036] The feedstock is a gasoline fraction containing sulfur compounds and olefins, whose boiling range generally extends from the boiling point of hydrocarbons having 2 or 3 carbon atoms (C2 or C3) to 260 °C, preferably from the boiling point of hydrocarbons having 2 or 3 carbon atoms (C2 or C3) to 220 °C, more preferably from the boiling point of hydrocarbons having 5 carbon atoms to 220 °C. The method according to the invention can also process feedstocks having an end point lower than those mentioned above, such as, for example, a C5-180 °C fraction.

[0037] The sulfur content of gasoline fractions produced by fluid catalytic cracking (FCC) depends on the sulfur content of the feedstock treated by said FCC, on the presence or absence of a pretreatment of the feedstock of the FCC, and on the end point of the fraction. Generally, the sulfur content of the whole gasoline fraction, especially those originating from FCC, is greater than 100 weight ppm, and most of the time greater than 500 weight ppm. For gasoline having an end point greater than 200 °C, the sulfur content is generally greater than 1000 weight ppm; in some cases, they can even reach values of about 4000 to 5000 weight ppm.

[0038] The feedstock to be treated by the method according to the invention can be a feedstock containing sulfur compounds with a sulfur content greater than 200 weight ppm and often greater than 500 ppm.

[0039] In addition, gasoline produced by a fluid catalytic cracking (FCC) unit contains on average from 0.5% to 5% by weight of diolefins, from 20% to 50% by weight of olefins and from 10 weight ppm to 0.5% by weight of sulfur, usually including less than 300 ppm of mercaptans.

[0040] Step a0): Selective hydrogenation step (optional)

[0041] Depending on the type of gasoline to be treated, it may be advantageous to pre-treat the aforementioned gasoline in the presence of hydrogen and a selective hydrogenation catalyst so as to at least partially hydrogenate the diolefins and to carry out a reaction to increase the molecular weight of a part of the light mercaptans (RSH) present in the feedstock by reaction with the olefins to obtain thioethers.

[0042] To this end, the gasoline to be processed is sent to a selective hydrogenation catalytic reactor, which contains at least one fixed bed or moving bed of a catalyst for the selective hydrogenation of diolefins and for increasing the molecular weight of light mercaptans. The reactions for the selective hydrogenation of diolefins and for increasing the molecular weight of light mercaptans are preferably carried out on a sulfided catalyst comprising at least one Group VIII metal and optionally at least one Group VIB metal and an oxide support. The Group VIII metal is preferably selected from nickel and cobalt, and especially nickel. The Group VIB metal, when present, is preferably selected from molybdenum and tungsten, and very preferably molybdenum.

[0043] The oxide support of the catalyst is preferably selected from alumina, nickel aluminate, silica, silicon carbide or a mixture of these oxides. Alumina is preferably used, and still more preferably high-purity alumina is used. According to a preferred embodiment, the selective hydrogenation catalyst contains nickel with a weight content of 1% to 12% of nickel oxide (in the form of NiO) and molybdenum with a weight content of 6% to 18% of molybdenum oxide (in the form of MoO3), and the nickel / molybdenum molar ratio is 0.3 to 2.5, and the metals are deposited on a support composed of alumina. The degree of sulfidation of the metals constituting the catalyst is preferably greater than 60%.

[0044] During the optional selective hydrogenation step, at a temperature of 50°C to 250°C, preferably 80°C to 220°C and still more preferably 90°C to 200°C, the gasoline is contacted with the catalyst at a space velocity (HSV) of 0.5 h -1 to 20 h -1 . The unit of the space velocity is the volume flow rate (l / l / h) of the raw material at 15°C per catalyst bed volume and per h. The pressure is 0.2 to 5 MPa, preferably 0.6 to 4 MPa and still more preferably 1 to 3 MPa. The optional selective hydrogenation step is generally carried out at a ratio of the hydrogen flow rate expressed as standard m 3 / m 3 of 2 to 100 Nm 3 / m 3 , preferably 3 to 30 Nm 3 / h to the volume flow rate of the raw material to be processed expressed as standard conditions (15°C, 0.1 MPa) of m 3 / h.

[0045] After the selective hydrogenation, the diolefin content determined via the maleic anhydride value (MAV) according to the UOP 326 method is generally reduced to less than 6 mg maleic anhydride / g, actually even less than 4 mg MA / g and more preferably less than 2 mg MA / g. In some cases, less than 1 mg MA / g can be obtained.

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

[0047] It should be noted that it is conceivable to simultaneously carry out the hydrogenation of the diolefins and the fractionation into two or three fractions by means of a catalytic distillation column, said catalytic distillation column comprising a distillation column equipped with at least one catalytic bed.

[0048] Step a): Selective hydrodesulfurization

[0049] The hydrodesulfurization step a) is carried out in order to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H2S.

[0050] The temperature is generally from 200 °C to 350 °C and preferably from 220 °C to 320 °C. The temperature used must be sufficient to keep the gasoline to be treated in the gas phase in the reactor.

[0051] The operating pressure of this step is generally from 0.2 MPa to 5 MPa and preferably from 1 MPa to 3 MPa.

[0052] The amount of catalyst used in each reactor of the first reaction section is generally such that the volume flow rate ratio of the gasoline to be treated in the catalyst bed at 15 °C (expressed as m 3 / h under standard conditions) (also known as the space velocity or HSV) is from 1 to 20 h 3 / h) (also known as the space velocity or HSV) is from 1 to 20 h -1 and preferably from 2 to 10 h -1 .

[0053] The hydrogen flow rate is generally such that the ratio of the hydrogen flow rate expressed as standard m 3 / h (Nm 3 / h) to the volume flow rate of the feedstock to be treated expressed as m 3 / h under standard conditions (15 °C, 0.1 MPa) is from 10 to 1000 Nm 3 / m 3 、preferably from 50 to 600 Nm 3 / m 3 . Standard m 3 is understood to mean the volume of 1 m 3 of gas at 0 °C and 0.1 MPa.

[0054] The hydrogen required for this step can be fresh hydrogen or recycled hydrogen, preferably H2S-free, or a mixture of fresh hydrogen and recycled hydrogen. Preferably, a mixture of fresh hydrogen and recycled hydrogen will be used.

[0055] Depending on the sulfur content of the feedstock to be treated, the degree of desulfurization in step a) is usually greater than 50% and preferably greater than 70%, such that the product resulting from step a) contains less than 200 ppm by weight of sulfur and preferably less than 100 ppm by weight of sulfur.

[0056] In the process according to the invention, the degree of hydrogenation of the olefins during this step a) is preferably less than 50%, more preferably less than 40%.

[0057] According to the invention, as described below, the hydrodesulfurization catalyst of step a) comprises an active phase and an oxide support, the active phase comprising at least one Group VIB metal and at least one Group VIII metal, optionally phosphorus, and preferably consisting of at least one Group VIB metal and at least one Group VIII metal, optionally phosphorus.

[0058] The Group VIB metal present in the active phase of the catalyst is preferably selected from molybdenum and tungsten.

[0059] The Group VIII metal present in the active phase of the catalyst is preferably selected from cobalt, nickel and mixtures of these two metals.

[0060] The active phase of the catalyst is preferably selected from the group formed by combinations of nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum, and very preferably the active phase consists of cobalt and molybdenum.

[0061] The content of the Group VIII metal, calculated as the oxide of the Group VIII metal, is preferably from 0.1% to 10% by weight relative to the total weight of the catalyst, more preferably from 0.6% to 8% by weight relative to the total weight of the catalyst, even more preferably from 0.6% to 7% by weight, and very preferably from 1% to 6% by weight. When the metal is cobalt or nickel, the metal content is expressed as CoO or NiO.

[0062] The total content of the Group VIB metal, calculated as the oxide of the Group VIB metal, is preferably from 1% to 20% by weight relative to the total weight of the catalyst, more preferably from 2% to 18% by weight relative to the total weight of the catalyst, and very preferably from 3% to 16% by weight. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 or WO3.

[0063] Preferably, the molar ratio of the Group VIII metal to the Group VIB metal of the catalyst is generally from 0.1 to 0.8 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0064] Optionally, relative to the total weight of the catalyst, the catalyst may additionally have a phosphorus content of P2O5 that is generally from 0.3 wt% to 10 wt%, preferably from 0.3 wt% to 5 wt%, and very preferably from 0.5 wt% to 3 wt%.

[0065] In addition, when phosphorus is present, the phosphorus / (Group VIB metal) molar ratio is generally from 0.1 to 0.7 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0066] Preferably, the catalyst of step a) has a specific surface area of 60 to 250 m 2 / g, preferably 60 to 200 m 2 / g and even more preferably 65 to 180 m 2 / g, and even more preferably 70 to 130 m 2 / g.

[0067] The total pore volume of the catalyst of step a) is generally from 0.3 cm 3 / g to 1.3 cm 3 / g, preferably from 0.4 cm 3 / g to 1.1 cm 3 / g.

[0068] The oxide support of the hydrodesulfurization catalyst is generally a porous solid selected from the group consisting of: alumina, silica, silica-alumina, or titanium oxide or magnesia used alone or in admixture with alumina or silica-alumina. It is preferably selected from the group consisting of silica, alumina, and silica-alumina. Very preferably, the oxide support consists essentially of alumina, that is, relative to the total weight of the support, it contains at least 51 wt%, preferably at least 60 wt%, very preferably at least 80 wt%, or even at least 90 wt% of alumina. It preferably consists only of alumina.

[0069] In a preferred embodiment, the catalyst of step a) comprises an alumina support and an active phase, the active phase comprising cobalt and molybdenum, preferably consisting of cobalt and molybdenum, the catalyst containing a weight content of cobalt oxide in the form of CoO of from 0.1 wt% to 10 wt%, preferably from 0.6 wt% to 8 wt%, more preferably from 0.6 wt% to 7 wt%, and even more preferably from 1 wt% to 6 wt% relative to the total weight of the catalyst, and a weight content of molybdenum oxide in the form of MoO3 of from 1 wt% to 20 wt%, preferably from 2 wt% to 18 wt%, and very preferably from 3 wt% to 16 wt% relative to the total weight of the catalyst, wherein the cobalt / molybdenum molar ratio is from 0.1 to 0.8 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0070] Preferably, the support of the hydrodesulfurization catalyst has a specific surface area of 60 to 250 m 2 / g, preferably 60 to 200 m 2 / g, and even more preferably 65 to 180 m 2 / g, and even more preferably 70 to 130 m 2 / g.

[0071] The total pore volume of the support of the hydrodesulfurization catalyst is generally 0.3 cm 3 / g to 1.3 cm 3 / g, preferably 0.4 cm 3 / g to 1.1 cm 3 / g.

[0072] The support of the hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometric shape, pellets, granules, compressed cylinders, crushed solids, or any other shaped body. Preferably, the support is in the form of beads having a diameter of 0.5 to 6 mm, or in the form of cylindrical, trilobal, or quadrilobal extrudates having an outside diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0073] Subsequently, the first partially desulfurized effluent obtained at the end of step a) is sent directly to step b) of the process according to the invention without separation.

[0074] Step b): Fine hydrodesulfurization (FNS) step

[0075] During the hydrodesulfurization step a), most of the sulfur compounds are converted to H2S. The remaining sulfur compounds are essentially refractory sulfur compounds, as well as the re - formed mercaptans resulting from the addition of the H2S formed in step a) to the olefins present in the feedstock.

[0076] Step b) of the process according to the invention comprises converting at least a portion of the re - formed mercaptans contained in the first effluent from step a) into olefins and H2S, and converting at least a portion of the sulfur compounds (such as thiophene compounds) contained in the first effluent from step a) into saturated compounds, for example into tetrahydrothiophene (or thiacyclopentane) or into mercaptans, and subsequently at least partially hydrogenolyzing these sulfur compounds to form H2S.

[0077] Preferably, step b) is carried out at a higher temperature than step a). Specifically, by using a higher temperature in this step compared to the temperature of step a), the formation of mercaptans is disadvantaged by changing the thermodynamic equilibrium. Step b) also enables the continued hydrodesulfurization of the residual sulfur compounds.

[0078] The temperature is usually from 250°C to 400°C, preferably from 270°C to 390°C. The temperature employed must be sufficient to keep the gasoline to be treated in the gas phase in the reactor.

[0079] The operating pressure at this stage is usually from 0.2 MPa to 5 MPa and preferably from 1.5 MPa to 3 MPa.

[0080] The amount of catalyst used in each reactor is usually such that the ratio (also known as the space velocity) expressed as the volume flow rate of the gasoline to be treated in m 3 at the standard conditions (15°C, 0.1 MPa) of the catalyst bed and expressed as m 3 / h is from 1 to 20 h -1 and preferably from 2 to 10 h -1 .

[0081] In the process according to the invention, during this step, the total olefin hydrogenation degree in step b) is preferably less than 30%.

[0082] Depending on the sulfur content of the feedstock to be treated, the total desulfurization degree in step b) is usually greater than 50% and preferably greater than 70%, such that the product produced in step b) contains less than 50 weight ppm of sulfur and preferably less than 20 weight ppm of sulfur, and even more preferably less than 10 weight ppm of sulfur.

[0083] The hydrofining hydrodesulfurization catalyst in step b) comprises an active phase and a porous alumina-based support, preferably consisting of an active phase and a porous alumina-based support. The active phase contains at least one Group VIB metal and at least one Group VIII metal and phosphorus, at least part of which is in the sulfide form. The content of the Group VIB metal in oxide form is from 1% to 20% by weight relative to the total weight of the catalyst, the content of the Group VIII metal in oxide form is from 0.2% to 10% by weight relative to the total weight of the catalyst, and the content of phosphorus in the form of P2O5 is from 0.1% to 5% by weight relative to the total weight of the catalyst. The catalyst has a specific surface area greater than or equal to 20 m 2 / g and less than 150 m 2 / g.

[0084] The content of the Group VIB metal in the active phase in oxide form is from 1% to 20% by weight, preferably from 2% to 15% by weight and more preferably from 3% to 12% by weight, relative to the total weight of the catalyst. The Group VIB metal is preferably molybdenum. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 or WO3.

[0085] The content of Group VIII metal in the active phase in terms of oxide form is 0.2 wt% to 10 wt%, preferably 0.5 wt% to 8 wt%, and even more preferably 0.5 wt% to 5 wt% relative to the total weight of the catalyst. The Group VIII metal is preferably cobalt or nickel. When the metal is cobalt, the cobalt content is expressed as CoO or NiO.

[0086] The content of phosphorus in terms of its P2O5 oxide form is 0.1 wt% to 5 wt%, preferably 0.2 wt% to 4 wt%, and even more preferably 0.3 wt% to 3 wt% relative to the total weight of the catalyst.

[0087] Preferably, the molar ratio of the Group VIII metal of the active phase to the Group VIB element of the active phase is 0.1 to 2.0 mol / mol, preferably 0.3 to 1.0 mol / mol.

[0088] Preferably, the molar ratio of phosphorus to the Group VIB metal of the active phase is 0.1 to 2.0 mol / mol, preferably 0.2 to 1.0 mol / mol, more preferably 0.2 to 0.5 mol / mol.

[0089] The specific surface area of the catalyst is greater than or equal to 20 m 2 / g and less than 150 m 2 / g, preferably 20 m 2 / g to 110 m 2 / g, and more preferably 25 m 2 / g to 90 m 2 / g.

[0090] The catalyst advantageously has a total pore volume measured by mercury porosimetry of 0.3 cm 3 / g to 1.3 cm 3 / g, preferably 0.35 cm 3 / g to 1.3 cm 3 / g, and very preferably 0.4 cm 3 / g to 1.2 cm 3 / g.

[0091] The support of the hydrofining hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometric shape, tablets, pellets, compressed cylinders, crushed solids or any other shaped body. Preferably, the support is in the form of beads with a diameter of 0.5 to 6 mm, or in the form of cylindrical, trilobal or quadrilobal extrudates with an outer diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0092] The carrier of the refined hydrodesulfurization catalyst contains alumina, preferably the carrier consists essentially of alumina, that is to say, relative to the total weight of the carrier, it contains at least 51% by weight of alumina, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight of alumina. Even more preferably, the carrier consists of alumina.

[0093] The specific surface area of the carrier is greater than or equal to 20 m 2 / g and less than 150 m 2 / g, preferably 20 m 2 / g to 110 m 2 / g, and more preferably 25 m 2 / g to 90 m 2 / g.

[0094] The carrier advantageously has a total pore volume measured by mercury porosimetry of 0.3 cm 3 / g to 1.3 cm 3 / g, preferably 0.35 cm 3 / g to 1.3 cm 3 / g, and very preferably 0.4 cm 3 / g to 1.2 cm 3 / g.

[0095] Step c): Separation of H2S [optional]

[0096] This step serves to separate the excess hydrogen and the H2S formed during steps a) and b). Any method known to those skilled in the art can be envisaged.

[0097] According to the first embodiment, after steps a) and b), the effluent is cooled to a temperature usually below 80 °C to condense the hydrocarbons. The gas phase and the liquid phase are then separated in a separation drum. The liquid fraction containing the desulfurized gasoline and a part of the dissolved H2S is sent to a stabilizer column or a debutanizer column. This column separates a top fraction consisting essentially of residual H2S and hydrocarbon compounds having a boiling point lower than or equal to the boiling point of butane and a bottoms fraction called stabilized gasoline containing compounds having a boiling point higher than the boiling point of n-butane from which H2S has been removed.

[0098] According to the second embodiment, after the condensation step, the liquid fraction containing the desulfurized gasoline and a part of the dissolved H2S is sent to a stripping section, while the gas fraction consisting mainly of hydrogen and H2S is sent to a purification section. Stripping can be carried out by heating the hydrocarbon fraction alone or in the presence of injected hydrogen or steam in a distillation column in order to extract at the top the light compounds entrained by dissolution in the liquid fraction and the dissolved residual H2S. The temperature of the stripped gasoline recovered at the bottom is usually from 120 °C to 250 °C.

[0099] Preferably, the separation step c) is carried out in a stabilizer column or a debutanizer column. This is because the stabilizer column enables more efficient separation of H2S than the stripping section.

[0100] Step c) is preferably carried out so that the sulfur in the form of H2S remaining in the desulfurized gasoline accounts for less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.

[0101] Preparation of catalyst

[0102] The catalysts used in steps a) and b) of the process according to the invention can be prepared by any technique known to those skilled in the art, in particular by impregnating a selected porous support with metals of Group VIII and Group VIB and optionally phosphorus. The impregnation can be carried out, for example, according to a method known to those skilled in the art under the term dry impregnation, in which exactly the required precursor amounts of the elements in the form of salts soluble in the selected solvent (such as softened water) are introduced so as to fill the pores of the support as precisely as possible. Preferably, when the impregnation aqueous solution contains cobalt, molybdenum and phosphorus, the impregnation aqueous solution is prepared under pH conditions that promote the formation of dissolved heteropolyanions. For example, the pH of such an aqueous solution is from 1 to 5. Preferably, the preparation of the catalyst is carried out without mixing and adding organic reagents with the precursors of the metals of Group VIII and Group VIB and phosphorus.

[0103] For example, in the molybdenum source, oxides and hydroxides, molybdic acid and its salts, especially ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PMo 12 O 40 ) and its salts, and optionally silicomolybdic acid (H4SiMo 12 O 40 ) and its salts can be used. The molybdenum source can also be any heteropoly compound of the Keggin, lacunary Keggin, substituted Keggin, Dawson, Anderson or Strandberg type, for example. Preferably, molybdenum trioxide and heteropoly compounds of the Keggin, lacunary Keggin, substituted Keggin and Strandberg types are used.

[0104] The cobalt precursors that can be used are advantageously selected from, for example, oxides, hydroxides, hydroxycarbonates, carbonates and nitrates. Cobalt hydroxide and cobalt carbonate are preferably used.

[0105] Phosphorus can be advantageously introduced alone or mixed with at least one of the Group VIB and Group VIII metals. Phosphorus is preferably introduced in admixture with the Group VIB and Group VIII metal precursors by dry impregnation of the porous support with a solution containing an elemental precursor and a phosphorus precursor. Preferred phosphorus sources are orthophosphoric acid H3PO4, but its salts and esters, such as ammonium phosphate or mixtures thereof, are also suitable. Phosphorus can also be introduced simultaneously with the Group VIB metal in the form of, for example, Keggin, lacunary Keggin, substituted Keggin or Strandberg type heteropolyanions.

[0106] The nickel precursors that can be used are advantageously selected from, for example, oxides, hydroxides, hydroxycarbonates, carbonates and nitrates. Nickel hydroxide and nickel hydroxycarbonate are preferably used.

[0107] The tungsten precursors that can be used are also known to those skilled in the art. For example, among tungsten sources, oxides and hydroxides, tungstic acid and its salts, especially ammonium salts, such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and its salts, and optionally silicotungstic acid (H4SiW 12 O 40 ) and its salts can be used. The tungsten source can also be any heteropoly compound of, for example, Keggin, lacunary Keggin, substituted Keggin or Dawson type. Oxides and ammonium salts, such as ammonium metatungstate, or heteropolyanions of Keggin, lacunary Keggin or substituted Keggin are preferably used.

[0108] The support filled with the solution can be aged at a temperature below 50 °C, preferably at ambient temperature, for a time not exceeding 12 h, preferably not exceeding 6 h.

[0109] After the aging step, the obtained catalyst precursor can be heat-treated. The purpose of this treatment is generally to convert the molecular precursors of the elements into oxide phases. In this case, this is an oxidation treatment, but simple drying of the catalyst can also be carried out.

[0110] In the case of drying, the catalyst precursor is dried at a temperature of 50 °C to below 200 °C, preferably 70 °C to 180 °C, for a period generally of 0.5 h to 12 h, and even more preferably for a period of 0.5 h to 5 h.

[0111] In the case of the oxidation treatment (also called calcination), the treatment is generally carried out under air or diluted oxygen, and the treatment temperature is generally 200 °C to 550 °C, preferably 300 °C to 500 °C, and advantageously lasts for a period generally of 0.5 h to 24 h, preferably for a period of 0.5 h to 12 h, and even more preferably for a period of 0.5 h to 10 h.

[0112] Before it is used as a hydrotreating catalyst, it is advantageous to subject the optionally dried or calcined catalyst to an activation step by sulfidation. This activation stage is carried out by methods known to those skilled in the art and is advantageously carried out in a sulfur-reducing atmosphere in the presence of hydrogen and hydrogen sulfide. Hydrogen sulfide can be used directly or generated from a sulfide reagent such as dimethyldisulfide.

[0113] Description of sulfidation of catalyst

[0114] Before contacting with the feedstock to be treated in a hydrodesulfurization process for gasoline, the catalyst used in the process according to the invention generally undergoes a sulfidation step. The sulfidation is preferably carried out in a sulfur-reducing medium, that is, in the presence of H2S and hydrogen, in order to convert metal oxides into sulfides such as MoS2, Co9S8 or Ni3S2. The sulfidation is carried out by injecting a stream containing H2S and hydrogen, or a sulfur compound and hydrogen capable of decomposing into H2S in the presence of the catalyst, into the catalyst. Polysulfides such as dimethyldisulfide (DMDS) are generally used as H2S precursors for sulfiding the catalyst. The sulfur can also be derived from the feedstock. The temperature is adjusted so that H2S reacts with the metal oxide to form metal sulfide. The sulfidation 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 a temperature of 200 to 600 °C and more preferably 300 to 500 °C.

[0115] The sulfidation degree of the metal constituting the catalyst is at least equal to 60%, preferably at least equal to 70%. The sulfur content in the sulfided catalyst is measured by elemental analysis according to ASTM D5373. When the overall sulfidation degree defined by the molar ratio of sulfur (S) present on the catalyst to the metal is at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of the relevant metal, the metal is considered to be sulfided. The total sulfidation degree is defined by the following equation:

[0116] (S / metal) 催化剂 ≥0.6×(S / metal) 理论

[0117] Where:

[0118] (S / metal) 催化剂 is the molar ratio of sulfur (S) present on the catalyst to the metal

[0119] (S / metal) 理论 is the molar ratio of sulfur to the metal corresponding to the complete sulfidation of the metal to obtain a sulfide.

[0120] This theoretical molar ratio varies according to the relevant metal:

[0121] -(S / Co)理论 = 1

[0122] -(S / Ni) 理论 = 1

[0123] -(S / Mo) 理论 = 2 / 1

[0124] -(S / W) 理论 = 2 / 1

[0125] When the catalyst contains multiple metals, the molar ratio of S to the combined metals present on the catalyst must also be at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of each metal to its sulfide, and this calculation is carried out in proportion to the relative molar fractions of each metal.

[0126] The following examples illustrate the invention without limiting its scope.

[0127] Examples (Examples without refining treatment)

[0128] The contents of Group VIII metals, Group VIB metals, and phosphorus were determined by X-ray fluorescence spectrometry.

[0129] The analytical methods used to characterize the feedstock and the effluent are as follows:

[0130] - For sulfur contents above 10 ppm S, the sulfur content according to ASTM D2622 method, and for sulfur contents below 10 ppm S, the sulfur content according to ISO 20846 method;

[0131] - The mercaptan content according to ASTM D3227 method;

[0132] - The olefin content based on gas chromatography analysis according to ASTM D6733 method.

[0133] Example 1: Preparation of catalyst A (according to the present invention)

[0134] A support A' composed of alumina in the form of beads is provided, which has a particle size of 2 to 4 mm and has a specific surface area of 35 m 2 / g and a pore volume of 0.57 ml / g.

[0135] Subsequently, cobalt, molybdenum, and phosphorus were added. The impregnation solution was prepared by dissolving molybdenum oxide (4.34 g, ≥99.5%, Sigma- ), cobalt hydroxide (1.03 g, 96%, Alfa ), and phosphoric acid (85 wt%, 0.95 g, 99.99%, Sigma- ) It is prepared by. After dry-impregnating 40 g of support A', the impregnated alumina is cured in a water-saturated atmosphere at ambient temperature for 4 h and then dried at 120 °C for 4 h. The catalyst thus obtained is designated as A.

[0136] Expressed in oxide form and reported relative to the weight of the dry catalyst, the final elemental composition of catalyst A is thus as follows: MoO3 = 9.5 ± 0.2 wt%, CoO = 1.8 ± 0.1 wt%, and P2O5 = 1.3 ± 0.1 wt%.

[0137] The Co / Mo and P / Mo molar ratios are 0.37 and 0.27, respectively.

[0138] The specific surface area of catalyst A is 31 m 2 / g.

[0139] Example 2: Preparation of catalyst B (according to the present invention)

[0140] There is provided a support B' composed of alumina in the form of cylindrical extrudates with a diameter of 1.6 mm, which has a specific surface area of 78 m 2 / g and a pore volume of 0.84 ml / g.

[0141] Subsequently, cobalt, molybdenum, and phosphorus are added. The impregnation solution is prepared by dissolving molybdenum oxide (3.67 g, ≥99.5%, Sigma- ) and cobalt hydroxide (0.87 g, 96%, Alfa ) and phosphoric acid (85 wt%, 0.73 g, 99.99%, Sigma- ) in 24 ml of softened water at 100 °C. After dry-impregnating 40 g of support B', the impregnated alumina is cured in a water-saturated atmosphere at ambient temperature for 4 h and then dried at 120 °C for 4 h. The catalyst thus obtained is designated as B.

[0142] Expressed in oxide form and relative to the weight of the dry catalyst, the final elemental composition of catalyst B is thus as follows: MoO3 = 8.2 ± 0.2 wt%, CoO = 1.5 ± 0.1 wt%, and P2O5 = 1.0 ± 0.1 wt%.

[0143] The Co / Mo and P / Mo molar ratios are 0.36 and 0.25, respectively.

[0144] The specific surface area of catalyst B is 75 m 2 / g.

[0145] Example 3: Preparation of catalyst C (not according to the present invention)

[0146] There is provided a support C' composed of alumina in the form of beads, which has a particle size of 2 to 4 mm and has a specific surface area of 139 m2 A specific surface area of / g and a pore volume of 0.97 ml / g.

[0147] Subsequently, nickel was added. The impregnation solution was prepared by dissolving nickel nitrate hexahydrate (34.36 g, ≥99.5%, Sigma- ) in 25 ml of softened water at room temperature. After dry impregnation of 40 g of support C’, the impregnated alumina was cured for 4 h in a water-saturated atmosphere at ambient temperature, then dried at 120 °C for 4 h, and finally calcined at 450 °C for 4 h under an air flow of 1 l / h / g. The catalyst thus obtained was designated as C.

[0148] Expressed in oxide form and reported relative to the weight of the dry catalyst, the final elemental composition of catalyst C was thus as follows: NiO = 17.9 ± 0.3 wt%.

[0149] The specific surface area of catalyst C was 114 m 2 / g.

[0150] Example 4: Preparation of catalyst D (not according to the present invention)

[0151] A support D’ composed of alumina in the form of beads was provided, which had a particle size of 2 to 4 mm and had a specific surface area of 194 m 2 / g and a pore volume of 0.60 ml / g.

[0152] Subsequently, cobalt, molybdenum and phosphorus were added. The impregnation solution was prepared by dissolving molybdenum oxide (2.24 g, ≥99.5%, Sigma- ), cobalt hydroxide (0.61 g, 96%, Alfa ) and phosphoric acid (85 wt%, 0.49 g, 99.99%, Sigma- ) in 17 ml of softened water at 100 °C. After dry impregnation of 40 g of support B’, the impregnated alumina was cured for 4 h in a water-saturated atmosphere at ambient temperature, then dried at 120 °C for 4 h. The catalyst thus obtained was designated as D.

[0153] Expressed in oxide form and reported relative to the weight of the dry catalyst, the final elemental composition of catalyst D was thus as follows: MoO3 = 5.2 ± 0.2 wt%, CoO = 1.1 ± 0.1 wt%, and P2O5 = 0.7 ± 0.1 wt%.

[0154] The Co / Mo and P / Mo molar ratios were 0.42 and 0.27, respectively.

[0155] The specific surface area of catalyst D was 189 m 2 / g.

[0156] Example 5: Preparation of catalyst E

[0157] Provide a support E' that is the same as support C'.

[0158] Subsequently, cobalt and molybdenum are added. The impregnation solution is prepared by dissolving ammonium heptamolybdate tetrahydrate (5.64 g, ≥99.5%, Sigma- ) and cobalt nitrate hexahydrate (5.36 g, ≥99.5%, Alfa ) in 28 ml of softened water at ambient temperature. After dry impregnation of 40 g of support E', the impregnated alumina is cured for 4 h in a water-saturated atmosphere at ambient temperature, then dried at 120 °C for 4 h, and finally calcined at 450 °C for 4 h under an air flow of 1 l / h / g. The catalyst thus obtained is designated as E.

[0159] Expressed in oxide form and reported relative to the weight of the dry catalyst, the final elemental composition of catalyst E is thus as follows: MoO3 = 10.0 ± 0.2 wt% and CoO = 3.0 ± 0.1 wt%.

[0160] The Co / Mo and P / Mo molar ratios are 0.60 and 0, respectively.

[0161] The specific surface area of catalyst E is 124 m 2 / g.

[0162] Example 6: Use of the catalyst in gasoline desulfurization process

[0163] Gasoline from the fluid catalytic cracking unit, consisting of 25 wt% olefins and 600 ppm S total sulfur, is subjected to a two-step treatment:

[0164] - The first step of selective hydrodesulfurization (HDS) using catalyst E in an adiabatic reactor. The operating conditions for the one-step hydrodesulfurization step of the gasoline feedstock are as follows: HSV = 3 h -1 , P = 2.0 MPa. A pure hydrogen stream is added to the feedstock at the reactor inlet such that H2 / HC = 250 Nm 3 / m 3 . The effluent is sent directly to the reactor of the second step;

[0165] - The second step of fine hydrodesulfurization (FNS) using catalysts A to D in an adiabatic reactor. Only the effluent from the first step is treated in this second step. The operating conditions for the refining step are as follows: HSV = 3 h -1 , P = 2.0 MPa. The reactor inlet temperature is always fixed 35 °C higher than the temperature of the effluent leaving the first selective hydrodesulfurization step.

[0166] Fix the inlet temperature of the reactor entering the first selective hydrodesulfurization step to obtain an effluent containing 10 weight ppm S of total sulfur (i.e., a total sulfur conversion rate of greater than 98%).

[0167] Before use, the catalyst in the selective and finishing hydrodesulfurization reactors is sulfided by contacting it with a feedstock consisting of sulfur in the form of 2 wt% dimethyldisulfide in n-heptane at 350 °C under a pressure of 3.4 MPa for 4 h.

[0168] The properties of the catalyst are given in Table 1. The performance properties in the gasoline desulfurization process are presented in Table 2.

[0169] The results show that, compared with using catalysts known in the prior art, the method for gasoline hydrodesulfurization according to the present invention can reduce the temperature level for treating gasoline feedstock in the first step of HDS, and can thus shorten the cycle time while maintaining a high level of desulfurization and limiting olefin loss, thereby maintaining product specifications.

[0170] Table 1

[0171]

[0172]

[0173] Table 2

[0174]

[0175] "Average temperature" corresponds to the weighted average bed temperature (WABT) well-known to those skilled in the art. The average temperature is advantageously determined as a function of the catalytic system used, the equipment items, and their configuration. The average temperature (or WABT) is calculated as follows:

[0176] WABT = (T 入口 + T 出口 ) / 2

[0177] where T 入口 : the temperature at the inlet of the reaction section, T 出口 : the temperature at the outlet of the reaction section. Unless otherwise specified, the "average temperature" of the reaction section is given under cycle start-up conditions.

Claims

1. A method for treating gasoline containing sulfur compounds and olefins, said method comprising at least the following steps: a) Contact gasoline, hydrogen, and a hydrodesulfurization catalyst at a temperature of 200 °C to 350 °C and a pressure of 0.2 MPa to 5 MPa, with a space velocity of 1 h -1 to 20 h -1 , expressed as the ratio of the hydrogen flow rate in standard m 3 / h to the volume flow rate of the feedstock to be treated in standard conditions, expressed as m 3 / h, is 10 Nm 3 / m 3 to 1000 Nm 3 / m 3 , and the hydrodesulfurization catalyst comprises an active phase and an oxide support, and the active phase comprises a Group VIB metal and a Group VIII metal at least partially in the form of sulfides, thereby obtaining a partially desulfurized effluent; b) Without separating the H2S formed in step a), directly contacting the partially desulfurized effluent obtained at the end of step a) with a refined hydrodesulfurization catalyst at a temperature of 250°C to 400°C and a pressure of 0.2 MPa to 5 MPa, where the space velocity is 1 h -1 to 20 h -1 , thereby obtaining a desulfurized effluent. The refined hydrodesulfurization catalyst comprises an active phase and a porous alumina-based support. The active phase preferably consists of at least one Group VIB metal and at least one Group VIII metal, at least part of which is in the sulfide form, and phosphorus. The content of the Group VIB metal in oxide form is 1 wt% to 20 wt% relative to the total weight of the catalyst, the content of the Group VIII metal in oxide form is 0.2 wt% to 10 wt% relative to the total weight of the catalyst, and the content of phosphorus in the form of P2O5 is 0.1 wt% to 5 wt% relative to the total weight of the catalyst. The catalyst has a specific surface area greater than or equal to 20 m 2 / g and less than 150 m 2 / g.

2. The method according to claim 1, wherein the molar ratio between the Group VIII metal of the active phase and the Group VIB metal of the active phase in the catalyst of step b) is from 0.1 to 2.0 mol / mol.

3. The method according to claim 1 or 2, wherein the molar ratio between phosphorus and the Group VIB metal of the active phase in the catalyst of step b) is from 0.1 to 2.0 mol / mol.

4. The method according to any one of the preceding claims, wherein the specific surface area of the catalyst in step b) is from 20 m 2 / g to 110 m 2 / g.

5. The method according to any one of the preceding claims, wherein the catalyst of step b) contains cobalt as the Group VIII metal and molybdenum as the Group VIB metal.

6. The method according to any one of the preceding claims, wherein the catalyst in step b) comprises an active phase composed of molybdenum, cobalt and phosphorus, and a porous alumina-based support, the cobalt content in the form of CoO oxide is 0.5 wt% to 5 wt% relative to the total weight of the catalyst, the molybdenum content in the form of MoO3 oxide is 3 wt% to 12 wt% relative to the total weight of the catalyst, the phosphorus content in the form of P2O5 oxide is 0.3 wt% to 3 wt% relative to the total weight of the catalyst, the molar ratio of cobalt to molybdenum is 0.3 to 1.0 mol / mol, the molar ratio of phosphorus to molybdenum is 0.2 to 0.5 mol / mol, and the specific surface area of the catalyst is 25 to 90 m 2 / g.

7. The method according to any one of the preceding claims, wherein the catalyst of step a) comprises an alumina support and an active phase containing cobalt and molybdenum, said catalyst containing a weight content of cobalt oxide in the form of CoO of from 0.1% to 10% by weight relative to the total weight of the catalyst, and a weight content of molybdenum oxide in the form of MoO3 of from 1% to 20% by weight relative to the total weight of the catalyst, wherein the cobalt / molybdenum molar ratio is from 0.1 to 0.8 mol / mol.

8. The method according to any one of the preceding claims, wherein the catalyst in step a) has a specific surface area of 60 to 250 m 2 / g.

9. The method according to any one of the preceding claims, wherein the temperature of step b) is higher than the temperature of step a).

10. The method according to any one of the preceding claims, wherein the gasoline is catalytic cracking gasoline.

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