FR3142487A1 - Hydrodesulfurization process for finishing gasolines using a catalyst based on group VIB and VIII metals and phosphorus on an alumina support with low specific surface area - Google Patents

Abstract

Process for treating a gasoline containing sulfur compounds and olefins, the process comprising at least the following steps: a) gasoline, hydrogen and a hydrodesulphurization catalyst comprising an active phase comprising a metal from group VIB and a metal from group VIII are brought into contact at least partly in sulphurized form, to obtain an effluent partially desulfurized; b) without separation of the H2S formed in step a), the partially desulfurized effluent obtained at the end of step a) is directly brought into contact with a catalyst comprising an active phase comprising at least one metal from group VIB and at least one metal from group VIII at least partly in sulphide form, phosphorus, and a porous support based on alumina, said catalyst having a surface area greater than or equal to 20 m²/g and less than 150 m2 /g.

Description

Process for the finishing hydrodesulfurization of gasolines using a catalyst based on metals from groups VIB and VIII and phosphorus on an alumina support with low specific surface area Field of invention

The present invention relates to the field of hydrotreatment of gasoline cuts, in particular gasoline cuts from fluidized bed catalytic cracking units. More particularly, the present invention relates to the use of catalysts in a process for producing low-sulfur gasoline. The invention applies particularly to the treatment of gasoline cuts containing olefins and sulfur, such as gasolines from catalytic cracking, for which it is sought to reduce the content of sulfur compounds, without hydrogenating the olefins and aromatics.

State of the art

The specifications for automotive fuels require a significant reduction in the sulphur content in these fuels, particularly in petrol. This reduction is intended to limit, in particular, the sulphur and nitrogen oxide content in automobile exhaust gases. The specifications currently in force in Europe since 2009 for petrol fuels set a maximum content of 10 ppm by weight (parts per million) of sulphur. Such specifications are also in force in other countries such as the United States and China, where the same maximum sulphur content has been required since January 2017. To achieve these specifications, it is necessary to treat petrol using desulphurisation processes.

The main sources of sulfur in gasoline bases are so-called cracked gasolines, and mainly the gasoline fraction resulting from a catalytic cracking process of a vacuum distillate or a residue from the atmospheric or vacuum distillation of crude oil. The gasoline fraction resulting from catalytic cracking, which represents on average 40% of gasoline bases, in fact contributes more than 90% of the sulfur content in gasolines. Consequently, the production of low-sulfur gasolines requires a desulfurization step for catalytic cracked gasolines. Other sources of gasoline that may contain sulfur include coker gasolines, visbreaker gasolines or, to a lesser extent, gasolines resulting from atmospheric distillation or steam cracking gasolines.

The removal of sulfur from gasoline cuts consists of specifically treating these sulfur-rich gasolines by desulfurization processes in the presence of hydrogen. These are called hydrodesulfurization (HDS) processes. However, these gasoline cuts and more particularly gasolines from fluidized catalytic cracking (or FCC for Fluid Catalytic Cracking according to the English terminology) contain a significant proportion of unsaturated compounds in the form of mono-olefins (approximately 20 to 50% by weight) which contribute to a good octane rating, diolefins (0.5 to 5% by weight) and aromatics. These unsaturated compounds are unstable and react during the hydrodesulfurization treatment. The diolefins form gums by polymerization during the hydrodesulfurization treatments. This formation of gums leads to a progressive deactivation of the hydrodesulfurization catalysts or a progressive clogging of the reactor. Consequently, diolefins must be removed by hydrogenation before any treatment of these gasolines. Traditional treatment processes desulfurize gasolines non-selectively by hydrogenating a large part of the mono-olefins, which results in a high loss in octane number and a high consumption of hydrogen. The most recent hydrodesulfurization processes make it possible to desulfurize cracked gasolines rich in mono-olefins, while limiting the hydrogenation of the mono-olefins and consequently the loss of octane. Such processes are for example described in documents EP-A-1077247 and EP-A-1174485.

However, in the case where cracked gasolines must be desulfurized very deeply, a portion of the olefins present in the cracked gasolines is hydrogenated on the one hand and recombines with H ₂ S to form mercaptans on the other hand. This family of compounds, with the chemical formula R-SH where R is an alkyl group, are generally called recombinant mercaptans, and generally represent between 20% by weight and 80% by weight of the residual sulfur in desulfurized gasolines. The reduction in the recombinant mercaptan content can be achieved by catalytic hydrodesulfurization, but this leads to the hydrogenation of a significant portion of the mono-olefins present in the gasoline, which then leads to a significant decrease in the octane number of the gasoline as well as an overconsumption of hydrogen. It is also known that the loss of octane linked to the hydrogenation of mono-olefins during the hydrodesulfurization stage is all the greater when the targeted sulfur content is low, i.e. when the aim is to eliminate in depth the sulfur compounds present in the feed.

It is thus possible to treat gasoline by a sequence of two reactors as described in document EP1077247, the first step, also called the selective HDS step, generally aims to achieve deep desulfurization of gasoline with a minimum of olefin saturation (and no aromatic loss) leading to maximum octane retention. The catalyst used is generally a CoMo type catalyst. During this step, new sulfur compounds are formed by recombination of the H ₂ S resulting from the desulfurization with the olefins: the recombinant mercaptans.

The second stage is generally used to minimize the amount of recombinant mercaptans. The temperature is generally higher in the second stage in order to thermodynamically promote the elimination of mercaptans. In practice, a furnace is therefore placed between the two reactors in order to be able to raise the temperature of the second reactor to a temperature higher than that of the first.

The catalyst used in the finishing process must be particularly selective so as not to induce saturation of olefins (and no aromatic loss) leading to a loss of octane. It must therefore make it possible to reduce the total sulfur and mercaptan contents of hydrocarbon cuts, preferably gasoline cuts, to very low contents, while minimizing the reduction in the octane number. Usually, the catalyst used is nickel-based.

It is known from patent FR 3,023,184 to propose a hydrotreatment catalyst on an alumina-based support comprising at least one metal from group VIB, one metal from group VII and phosphorus, having a specific surface area of between 20 and 150 m ² /g.

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

However, there is still a need to maximize performance in the hydrotreatment of gasoline cuts to achieve sulfur specifications.

Surprisingly, the applicant has identified that a catalyst based on at least one metal from group VIII, at least one metal from group VIB and phosphorus on an alumina support with a low specific surface area within a range of well-defined values, when implemented in the finishing hydrodesulfurization (FNS) section downstream of the selective hydrodesulfurization (HDS) section, allows a reduction in the treatment temperatures in gasoline hydrodesulfurization while maintaining the performance of the process in terms of reducing the sulfur content and preserving the olefins to maintain the octane number. A reduction in the average treatment temperature in the HDS section makes it possible to increase the overall cycle time of the process.

Without being tied to any theory, the use of a type of finishing hydrodesulfurization catalyst makes it possible to eliminate part of the refractory sulfur compounds in the finishing section while preserving the olefins induced by a high selectivity permitted by a specific interaction between the active phase and the surface of the alumina support with low specific surface area, thus overall on the selective hydrodesulfurization and finishing sections to better control the hydrogenation reactions of the olefins and conversion of the sulfur compounds while minimizing the treatment temperature.

Objects of the invention

The purpose of the pre-sale invention is to implement a process for producing low-sulfur gasolines, making it possible to valorize the entire 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 reducing the gasoline yield while minimizing the reduction in the octane number due to the hydrogenation of olefins.

The present invention relates to a process for treating gasoline containing sulfur compounds and olefins, the process comprising at least the following steps:

a) bringing into contact gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a metal from group VIB and a metal from group VIII at least partly in sulphide form, and an oxide support, at a temperature of between 200°C and 350°C, at a pressure of between 0.2 MPa and 5 MPa, with a space velocity of between 1 h ^-1 and 20 h ^-1 and a ratio between the hydrogen flow rate expressed in normal m ³ per hour and the flow rate of feed to be treated expressed in m ³ per hour under standard conditions of between 10 Nm ³ /m ³ and 1000 Nm ³ /m ³ , to obtain a partially desulphurized effluent;

b) without separation of the H ₂ S formed in step a), the partially desulfurized effluent obtained at the end of step a) is brought directly into contact with a finishing hydrodesulfurization catalyst at a temperature of between 250°C and 400°C, at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 and 20 h ^-1 , to obtain a desulfurized effluent, said finishing hydrodesulfurization catalyst comprising an active phase comprising, preferably consisting of, at least one metal from group VIB and at least one metal from group VIII at least partly in sulfurized form, phosphorus, and a porous support based on alumina, the content of metal from group VIB, measured in oxide form, being between 1 and 20% by weight relative to the total weight of the catalyst, the content of metal from group VIII, measured in oxide form, being between 0.2 and 10% by weight relative to to the total weight of the catalyst, and the phosphorus content, measured in its P ₂ O ₅ form, being between 0.1 and 5% by weight relative to the total weight of the catalyst, said catalyst having a surface area greater than or equal to 20 m²/g and less than 150 m ² /g.

According to one or more embodiments, the catalyst of step b) comprises a molar ratio between said group VIII metal of the active phase and said group VIB metal of the active phase between 0.1 and 2.0 mol/mol.

According to one or more embodiments, the catalyst of step b) comprises a molar ratio between phosphorus and said group VIB metal of the active phase between 0.1 and 2.0 mol/mol.

According to one or more embodiments, the catalyst of step b) comprises a specific surface area between 20 m ² /g and 110 m ² /g.

According to one or more embodiments, the catalyst of step b) comprises cobalt as a group VIII metal and molybdenum as a group VIB metal.

According to one or more embodiments, wherein the catalyst of step b) comprises an active phase consisting of molybdenum, cobalt and phosphorus, and a porous support based on alumina, the cobalt content being between 0.5 and 5% by weight, measured in the oxide form CoO, relative to the total weight of the catalyst, the molybdenum content being between 3 and 12% by weight, measured in the oxide form MoO ₃ , relative to the total weight of the catalyst, a phosphorus content of between 0.3 and 3% by weight, measured in the oxide form P ₂ O ₅ , relative to the total weight of the catalyst, the molar ratio between cobalt and molybdenum being between 0.3 and 1.0 mol/mol, the molar ratio between phosphorus and molybdenum being between 0.2 and 0.5 mol/mol, the specific surface area of the catalyst being between 25 ^m2 /g and 90 ^m2 /g.

According to one or more embodiments, the catalyst of step a) comprises an alumina support and an active phase comprising cobalt and molybdenum, 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% by weight, and a content by weight relative to the total weight of catalyst of molybdenum oxide, in MoO ₃ form, of between 1 and 20% by weight, with a cobalt/molybdenum molar ratio of between 0.1 and 0.8 mol/mol.

According to one or more embodiments, the catalyst of step a) comprises a specific surface area of between 60 and 250 m²/g.

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

According to one or more embodiments, the gasoline is a catalytic cracking gasoline.

Detailed description Definitions

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 D.R. Lide, 81st edition, 2000-2001). For example, group VIII (or VIIIB) according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification.

The BET specific surface area is measured by nitrogen physisorption according to ASTM D3663-03, a method described in Rouquerol F.; Rouquerol J.; Singh K. “Adsorption by Powders & Porous Solids: Principle, methodology and applications”, Academic Press, 1999.

In the following description of the invention, the total pore volume of the oxide support or catalyst is understood to mean the volume measured by intrusion with a mercury porosimeter according to ASTM D4284-83 at a maximum pressure of 4000 bar (400 MPa), using a surface tension of 484 dyne/cm and a contact angle of 140°. The wetting angle was taken to be equal to 140° following the recommendations of the work “Techniques de l'ingénieur, traité analyse et caractérisation”, p.1050-5, written by Jean Charpin and Bernard Rasneur.

For greater accuracy, the value of the total pore volume in ml/g or ^cm3 /g given in the following text corresponds to the value of the total mercury volume (total pore volume measured by intrusion with a mercury porosimeter) in ml/g or ^cm3 /g measured on the sample minus the value of the mercury volume in ml/g or ^cm3 /g measured on the same sample for a pressure corresponding to 30 psi (approximately 0.2 MPa).

The contents of group VIII metal, group VIB metal and phosphorus are measured by X-ray fluorescence.

The contents of group VIB metals, group VIII metals, and phosphorus in the catalyst are expressed as oxides after correction for the loss on ignition of the catalyst sample at 550°C for two hours in a muffle furnace. The loss on ignition is due to moisture loss. It is determined according to ASTM D7348.

The charge

The process according to the invention makes it possible to treat any type of gasoline cut containing sulfur compounds and olefins, alone or in a mixture, such as, for example, a cut from a coking unit, visbreaking unit, steam cracking unit or fluid catalytic cracking unit (FCC). This gasoline may optionally be composed of a significant fraction of gasoline from other production processes such as atmospheric distillation (straight run gasoline) or conversion processes (coking or steam cracking gasoline). Said feedstock preferably consists of a gasoline cut from a catalytic cracking unit.

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.

The sulfur content of gasoline cuts produced by catalyst cracking (FCC) depends on the sulfur content of the feedstock treated by the FCC, the presence or absence of a pretreatment of the FCC feedstock, and the end point of the cut. Generally, the sulfur contents of an entire gasoline cut, particularly those from FCC, are greater than 100 ppm by weight and most of the time greater than 500 ppm by weight. For gasolines with 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.

The feedstock treated by the process according to the invention may be a feedstock containing sulfur compounds in a content greater than 200 ppm by weight of sulfur, and often greater than 500 ppm.

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 a0) Selective hydrogenation (optional)

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 increase the weight of part of the light mercaptans (RSH) present in the charge in thioethers, by reaction with olefins.

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 metal from group VIII and optionally at least one metal from group VIB and an oxide support. The metal from group VIII is preferably chosen from nickel and cobalt and in particular nickel. The metal from group VIB, when present, is preferably chosen from molybdenum and tungsten and very preferably molybdenum.

The oxide support of the catalyst is preferably chosen from alumina, nickel aluminate, silica, silicon carbide, or a mixture of these oxides. Alumina is preferably used and even more preferably high-purity alumina. According to a preferred embodiment, the selective hydrogenation catalyst contains nickel at a content by weight of nickel oxide, in the form of NiO, of between 1 and 12%, and molybdenum at a content by weight of molybdenum oxide, in the form of MoO ₃ , 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%.

In the optional selective hydrogenation step, the gasoline is contacted with the catalyst at a temperature of between 50 and 250°C, and preferably between 80 and 220°C, and even more preferably between 90 and 200°C, with an hourly volumetric flow rate (HVV) of between 0.5 h ^-1 and 20 h ^-1 , the unit of the hourly volumetric flow rate being the volume flow rate of charge at 15°C per volume of catalytic bed and per hour (L/L/h). The pressure is between 0.2 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 volume flow rate of feedstock to be treated expressed in ^m3 per hour at standard conditions (15°C, 0.1 MPa) of between 2 and 100 ^Nm3 / ^m3 , preferably between 3 and 30 ^Nm3 / ^m3 .

After selective hydrogenation, the diolefin content, determined by means of the maleic anhydride value (MAV), according to the UOP 326 method, is generally reduced to less than 6 mg maleic anhydride/g, or even less than 4 mg AM/g and more preferably less than 2 mg AM/g. In some cases, less than 1 mg AM/g may be obtained.

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.

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) Selective hydrodesulfurization

The hydrodesulfurization step a) is implemented to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H ₂ S.

The temperature is generally between 200°C and 350°C, and preferably between 220°C and 320°C. The temperature used must be sufficient to maintain the gasoline to be treated in the gas phase in the reactor.

The operating pressure of this step is generally between 0.2 MPa and 5 MPa, preferably between 1 MPa and 3 MPa.

The quantity of catalyst used in each reactor of the first reaction section is generally such that the ratio between the volume flow rate at 15°C of gasoline to be treated expressed in ^m3 per hour at standard conditions, per ^m3 of catalytic bed (also called hourly volumetric velocity - HVV) is between 1 and 20 h ^-1 and preferably between 2 and 10 h ^-1 .

The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m ³ per hour (Nm ³ /h) and the volume flow rate of the feedstock to be treated expressed in m ³ per hour at standard conditions (15°C, 0.1 MPa) is between 10 and 1000 Nm ³ /m ³ , preferably between 50 and 600 Nm ³ /m ³ . Normal m ³ means the volume of 1 m ³ of gas at 0°C and 0.1 MPa.

The hydrogen required for this step may be fresh hydrogen or recycled hydrogen, preferably freed from H ₂ S, or a mixture of fresh hydrogen and recycled hydrogen. Preferably, a mixture of fresh hydrogen and recycled hydrogen will be used.

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 200 ppm by weight of sulfur and preferably less than 100 ppm by weight of sulfur.

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

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

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 metals.

The active phase of the catalyst is preferably selected from the group formed by the combination of the metals nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum and very preferably the active phase consists of cobalt and molybdenum.

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

The group VIB metal content is preferably between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst, more preferably between 2 and 18% by weight, and very preferably between 3 and 16% by weight of group VIB metal oxide relative to the total weight of the catalyst. When the metal is molybdenum or tungsten, the metal content is expressed as MoO ₃ or WO ₃ .

Preferably, the molar ratio of group VIII metal to group VIB metal of the catalyst is generally between 0.1 and 0.8 mol/mol, preferably between 0.2 and 0.6 mol/mol.

Optionally, the catalyst may further have a phosphorus content generally between 0.3 and 10% by weight of P ₂ O ₅ relative to the total weight of catalyst, preferably between 0.3 and 5% by weight, very preferably between 0.5 and 3% by weight.

Furthermore, when phosphorus is present, the phosphorus/(group VIB metal) molar ratio is generally between 0.1 and 0.7 mol/mol, preferably between 0.2 and 0.6 mol/mol.

Preferably, the catalyst of step a) comprises a specific surface area of between 60 and 250 m²/g, preferably of between 60 and 200 m²/g, and even more preferably of between 65 and 180 m²/g, and even more preferably of between 70 and 130 m²/g.

The total pore volume of the catalyst of step a) is generally between 0.3 cm ³ /g and 1.3 cm ³ /g, preferably between 0.4 cm ³ /g and 1.1 cm ³ /g.

The oxide support of the hydrodesulfurization catalyst is typically a porous solid selected from the group consisting of: aluminas, silica, silica-alumina or titanium or magnesium oxides used alone or in a mixture with alumina or silica-alumina. It is preferably selected from the group consisting of silica, alumina, and silica-alumina. Very preferably, the oxide support is essentially composed 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 relative to the total weight of said oxide support. It is preferably composed solely of alumina.

In a preferred embodiment, the catalyst of step a) comprises an alumina support and an active phase comprising, preferably consisting of, cobalt and molybdenum, 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% by weight, preferably between 0.6 and 8% by weight, more preferably between 0.6 and 7% by weight and even more preferably between 1 and 6% by weight, and a content by weight relative to the total weight of catalyst of molybdenum oxide, in MoO ₃ form, of between 1 and 20% by weight, preferably between 2 and 18% by weight, and very preferably between 3 and 16% by weight, with a cobalt/molybdenum molar ratio of between 0.1 and 0.8 mol/mol, preferably between 0.2 and 0.6 mol/mol.

Preferably, the support of the hydrodesulfurization catalyst comprises a specific surface area of between 60 and 250 m²/g, preferably of between 60 and 200 m²/g, and even more preferably of between 65 and 180 m²/g, and even more preferably of between 70 and 130 m²/g.

The total pore volume of the hydrodesulfurization catalyst support is generally between 0.3 cm ³ /g and 1.3 cm ³ /g, preferably between 0.4 cm ³ /g and 1.1 cm ³ /g.

The support of the hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometry, platelets, pellets, compressed cylinders, crushed solids or any other shape. 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 a circumscribed diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

The first 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) Finishing hydrodesulfurization step (FNS)

During the hydrodesulfurization step a), a large part of the sulfur compounds are transformed into H ₂ S. The remaining sulfur compounds are essentially refractory sulfur compounds and the recombinant mercaptans resulting from the addition of the H ₂ S formed in step a) to the olefins present in the feed.

Step b) of the process according to the invention consists in transforming at least part of the recombinant mercaptans contained in the first effluent from step a) into olefins and H ₂ S as well as at least part of the sulfur compounds contained in the first effluent from step a) such as thiophene compounds, into saturated compounds for example thiophanes (or thiacyclopentanes) or mercaptans, then in at least partially hydrogenolyzing these sulfur compounds to form H ₂ S.

Preferably, step b) is carried out at a higher temperature than that of step a). Indeed, by using a higher temperature in this step compared to the temperature of step a), the formation of mercaptans will be disadvantaged by displacement of the thermodynamic equilibrium. Step b) also makes it possible to continue the hydrodesulfurization of the residual sulfur compounds.

The temperature is generally between 250°C and 400°C, preferably between 270°C and 390°C. The temperature used must be sufficient to maintain the gasoline to be treated in the gas phase in the reactor.

The operating pressure of this step is generally between 0.2 MPa and 5 MPa and preferably between 1.5 MPa and 3 MPa.

The quantity of catalyst used in each reactor is generally such that the ratio between the volume flow rate of gasoline to be treated expressed in ^m3 per hour at standard conditions (15°C, 0.1 MPa), per ^m3 of catalytic bed (also called hourly volume flow rate) is between 1 and 20 h ^-1 and preferably between 2 and 10 h ^-1 .

In the process according to the invention, the hydrogenation rate of the olefins in step b) is preferably less than 30%.

The total desulfurization rate of step b), which depends on the sulfur content of the feedstock to be treated, is generally greater than 50% and preferably greater than 70% so that the product from step b) contains less than 50 ppm by weight of sulfur and preferably less than 20 ppm by weight of sulfur, and even more preferably less than 10 ppm by weight of sulfur.

The finishing hydrodesulfurization catalyst of step b) preferably comprises, and consists of, an active phase containing at least one metal from group VIB and at least one metal from group VIII at least partly in sulfurized form, phosphorus, and a porous support based on alumina, the content of metal from group VIB, measured in oxide form, is between 1 and 20% by weight relative to the total weight of the catalyst, the content of metal from group VIII, measured in oxide form, is between 0.2 and 10% by weight relative to the total weight of the catalyst, and the content of phosphorus, measured in its oxide form P ₂ O ₅ , being between 0.1 and 5% by weight relative to the total weight of the catalyst, said catalyst having a specific surface area greater than or equal to 20 m²/g and less than 150 m²/g.

The group VIB metal content of the active phase, measured in oxide form, is between 1 and 20% by weight relative to the total weight of the catalyst, preferably between 2 and 15% by weight, and even more preferably between 3 and 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 MoO ₃ or WO ₃ .

The group VIII metal content of the active phase, measured in oxide form, is between 0.2 and 10% by weight relative to the total weight of the catalyst, preferably between 0.5 and 8% by weight, and even more preferably between 0.5 and 5% by weight relative to the total weight of the catalyst. The group VIII metal is preferably cobalt. When the metal is cobalt or nickel, the cobalt content is expressed as CoO or NiO.

The phosphorus content, measured in its oxide form P ₂ O ₅ , is between 0.1 and 5% by weight relative to the total weight of the catalyst, preferably between 0.2 and 4% by weight, and even more preferably between 0.3 and 3% by weight relative to the total weight of the catalyst.

Preferably, the molar ratio between said group VIII metal of the active phase and said group VIB metal of the active phase is between 0.1 and 2.0 mol/mol, preferably between 0.3 and 1.0 mol/mol.

Preferably, the molar ratio between phosphorus and said group VIB metal of the active phase is between 0.1 and 2.0 mol/mol, preferably between 0.2 and 1.0 mol/mol, more preferably between 0.2 and 0.5 mol/mol.

The specific surface area of the catalyst is greater than or equal to 20 m²/g and less than 150 ^m2 /g, preferably between 20 ^m2 /g and 110 ^m2 /g, more preferably between 25 ^m2 /g and 90 ^m2 /g.

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

The support of the finishing hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometry, platelets, pellets, compressed cylinders, crushed solids or any other shape. 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 a circumscribed diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

The support of said finishing hydrodesulfurization catalyst comprises alumina, preferably the support is mainly composed 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 relative to the total weight of said support. It is preferably composed solely of alumina.

The specific surface area of the support is greater than or equal to 20 m²/g and less than 150 ^m2 /g, preferably between 20 ^m2 /g and 110 ^m2 /g, more preferably between 25 ^m2 /g and 90 ^m2 /g.

The support advantageously has a total pore volume measured by mercury porosimetry of between 0.3 cm ³ /g and 1.3 cm ³ /g, preferably between 0.35 cm ³ /g and 1.3 cm ³ /g, and very preferably between 0.4 cm ³ /g and 1.2 cm ³ /g.

Step c): Separation of H2S [optional]

This step is implemented in order to separate the excess hydrogen as well as the H ₂ S formed during steps a) and b). Any method known to those skilled in the art may be considered.

According to a first embodiment, after steps a) and b), the effluent is cooled to a temperature generally below 80°C in order to condense the hydrocarbons. The gas and liquid phases are then separated in a separation tank. The liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved H ₂ S is sent to a stabilization column or debutanizer. This column separates a top cut essentially consisting of residual H ₂ S and hydrocarbon compounds having a boiling point less than or equal to that of butane and a bottom cut freed from H ₂ S, called stabilized gasoline, containing the compounds having a boiling point higher than that of n-butane.

According to a second embodiment, after the condensation step, the liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved H ₂ S is sent to a stripping section, while the gaseous fraction consisting mainly of hydrogen and H ₂ S is sent to a purification section. The stripping can be carried out by heating the hydrocarbon fraction alone or with an injection of hydrogen or water vapor, in a distillation column in order to extract at the top, the light compounds which have been entrained by dissolution in the liquid fraction as well as the dissolved residual H ₂ S. The temperature of the stripped gasoline recovered at the bottom of the column is generally between 120°C and 250°C.

Preferably, separation step c) is carried out in a stabilization column or debutanizer. Indeed, a stabilization column allows H ₂ S to be separated more efficiently than a stripping section.

Step c) is preferably carried out so that the sulfur in the form of H ₂ S remaining in the desulfurized gasoline represents less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.

Preparation of catalysts

The catalysts used in steps a) and b) of the process according to the invention can be prepared using any technique known to those skilled in the art, and in particular by impregnation of the metals of groups VIII and VIB and optionally phosphorus on the selected porous support. The impregnation can for example be carried out according to the method known to those skilled in the art under the terminology of dry impregnation, in which just the amount of precursors of desired elements in the form of salts soluble in the chosen solvent, for example demineralized water, is introduced so as to fill the porosity of the support as exactly as possible. Preferably, the aqueous impregnation solution when it contains cobalt, molybdenum and phosphorus is prepared under pH conditions promoting the formation of heteropolyanions in solution. For example, the pH of such an aqueous solution is between 1 and 5. Preferably, the preparation of the catalyst is carried out without adding organic agent in a mixture with the precursors of the metals of group VIII, group VI and phosphorus.

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 (H ₃ PMo ₁₂ O ₄₀ ), and their salts, and possibly silicomolybdic acid (H ₄ SiMo ₁₂ O ₄₀ ) 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. Preferably, molybdenum trioxide and heteropolycompounds of the Keggin, lacunar Keggin, substituted Keggin and Strandberg type are used.

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

Phosphorus may advantageously be introduced alone or in admixture with at least one of the metals of group VIB and VIII. The phosphorus is preferably introduced in admixture with the precursors of the metals of group VIB and group VIII by dry impregnation of said porous support using a solution containing the precursors of the elements and the precursor of phosphorus. The preferred source of phosphorus is orthophosphoric acid H ₃ PO ₄ , but its salts and esters such as ammonium phosphates or their mixtures are also suitable. The phosphorus may also be introduced at the same time as the metal(s) of group VIB in the form of, for example, Keggin, Keggin lacunar, Keggin substituted or Strandberg type heteropolyanions.

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

The tungsten precursors that can be used are also well known to those skilled in the art. For example, among the tungsten sources, 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 optionally silicotungstic acid (H ₄ SiW ₁₂ O ₄₀ ) and its salts. The tungsten sources can also be any heteropolycompound of the Keggin, Keggin lacunary, Keggin substituted, Dawson type, for example. Preferably, ammonium oxides and salts such as ammonium metatungstate or heteropolyanions of the Keggin, Keggin lacunary or Keggin substituted type are used.

The support thus filled with the solution can be left to mature at a temperature below 50°C, preferably at room temperature, for a time not exceeding 12 hours, preferably not exceeding 6 hours.

Following the maturation stage, the catalyst precursor obtained can undergo a heat treatment. This treatment generally aims to transform the molecular precursors of the elements into the oxide phase. In this case, it is an oxidizing treatment but a simple drying of the catalyst can also be carried out.

In the case of drying, the catalyst precursor is dried at a temperature of between 50°C and less than 200°C, preferably between 70°C and 180°C, for a period of time typically of between 0.5 hours and 12 hours, and even more preferably for a period of time of between 0.5 hours and 5 hours.

In the case of an oxidizing treatment, also called calcination, this is generally carried out in air or in diluted oxygen, and the treatment temperature is generally between 200°C and 550°C, preferably between 300°C and 500°C, and advantageously for a duration typically between 0.5 hours and 24 hours, preferably for a duration of 0.5 hours to 12 hours, and even more preferably for a duration of 0.5 hours to 10 hours.

Before its use as a hydrotreatment catalyst, it is advantageous to subject the optionally dried or calcined catalyst to an activation step by sulfurization. This activation phase is carried out by methods well known to those skilled in the art, and advantageously under a sulfide-reducing atmosphere in the presence of hydrogen and hydrogen sulfide. The hydrogen sulfide can be used directly or generated by a sulfide agent (such as dimethyl disulfide).

Description of catalyst sulfidation

Before contacting with the feedstock to be treated in a gasoline hydrodesulfurization process, the catalysts used in the process according to the invention generally undergo a sulfurization step. The sulfurization is preferably carried out in a sulforeducing medium, i.e. in the presence of H ₂ S and hydrogen, in order to transform the metal oxides into sulfides such as, for example, MoS ₂ , Co ₉ S ₈ or Ni ₃ S ₂ . The sulfurization is carried out by injecting onto the catalyst a flow containing H ₂ S and hydrogen, or a sulfur compound capable of decomposing into H ₂ S in the presence of the catalyst and hydrogen. Polysulfides such as dimethyl disulfide (DMDS) are H ₂ S 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.

The sulfurization rate of the metals constituting the catalysts is at least equal to 60%, preferably at least equal to 70%. 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. The overall sulfurization rate is defined by the following equation:

(S/metal) _catalyst ≥ 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/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.

The following examples illustrate the invention without limiting its scope.

Examples (example without finishing)

The contents of group VIII, group VIB metals and phosphorus are measured by X-ray fluorescence.

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

- sulfur content according to ASTM D2622 method for contents greater than 10 ppm S and ISO 20846 for contents less than 10 ppm S;

- mercaptan content according to ASTM D3227 method;

- olefin content based on gas chromatographic analysis according to ASTM D6733 method.

Example 1: Preparation of catalyst A (compliant)

A support A' composed of alumina in the form of beads with a particle size of between 2 and 4 mm, and having a specific surface area of 35 m ² /g and a pore volume of 0.57 mL/g is provided.

Cobalt, molybdenum and phosphorus are then added. The impregnation solution is prepared by dissolving at 100°C molybdenum oxide (4.34 g, ≥99.5%, Sigma-Aldrich®), cobalt hydroxide (1.03 g, 96%, Alfa Aesar®), 85% phosphoric acid by weight (0.95 g, 99.99%, Sigma-Aldrich®) in 16 mL of demineralized water. After dry impregnation of 40 grams of support A’, the impregnated alumina is left to mature in a water-saturated atmosphere for 4 hours at room temperature, then dried at 120°C for 4 hours. The catalyst thus obtained is denoted A.

The final composition of elements of catalyst A, expressed in the form of oxides and relative to the weight of the dry catalyst is then as follows: MoO ₃ = 9.5 +/- 0.2% by weight, CoO = 1.8 +/- 0.1% by weight and P ₂ O ₅ = 1.3 +/- 0.1% by weight.

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

The specific surface area of catalyst A is 31 ^m2 /g.

Example 2: Preparation of catalyst B (compliant)

A support B' composed of alumina is provided in the form of cylindrical extrudates of 1.6 mm diameter, and having a specific surface area of 78 m ² /g and a pore volume of 0.84 mL/g.

Cobalt, molybdenum and phosphorus are then added. The impregnation solution is prepared by dissolving at 100°C molybdenum oxide (3.67 g, ≥99.5%, Sigma-Aldrich®), cobalt hydroxide (0.87 g, 96%, Alfa Aesar®), 85% phosphoric acid by weight (0.73 g, 99.99%, Sigma-Aldrich®) in 24 mL of demineralized water. After dry impregnation of 40 grams of support B’, the impregnated alumina is left to mature in a water-saturated atmosphere for 4 hours at room temperature, then dried at 120°C for 4 hours. The catalyst thus obtained is denoted B.

The final composition of elements of catalyst B, expressed in the form of oxides and relative to the weight of the dry catalyst is then as follows: MoO ₃ = 8.2 +/- 0.2% by weight, CoO = 1.5 +/- 0.1% by weight and P ₂ O ₅ = 1.0 +/- 0.1% by weight.

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

The specific surface area of catalyst B is 75 ^m2 /g.

Example 3: Preparation of catalyst C (non-compliant)

A support C' composed of alumina in the form of beads with a particle size of between 2 and 4 mm, and having a specific surface area of 139 m ² /g and a pore volume of 0.97 mL/g, is provided.

Nickel is then added. The impregnation solution is prepared by dissolving nickel nitrate hexahydrate (34.36 g, ≥99.5%, Sigma-Aldrich®) in 25 mL of demineralized water at room temperature. After dry impregnation of 40 grams of support C’, the impregnated alumina is left to mature in a water-saturated atmosphere for 4 hours at room temperature, then dried at 120°C for 4 hours, and finally calcined under an air flow rate of 1 L/h/g at 450°C for 4 hours. The catalyst thus obtained is denoted C.

The final composition of elements in catalyst C, expressed in the form of oxides and relative to the weight of the dry catalyst, is then as follows: NiO = 17.9 +/- 0.3% by weight.

The specific surface area of catalyst C is 114 ^m2 /g.

Example 4: Preparation of catalyst D (non-compliant)

A support D' composed of alumina in the form of beads with a particle size of between 2 and 4 mm, and having a specific surface area of 194 m ² /g and a pore volume of 0.60 mL/g is provided.

Cobalt, molybdenum and phosphorus are then added. The impregnation solution is prepared by dissolving at 100°C molybdenum oxide (2.24 g, ≥99.5%, Sigma-Aldrich®), cobalt hydroxide (0.61 g, 96%, Alfa Aesar®), 85% phosphoric acid by weight (0.49 g, 99.99%, Sigma-Aldrich®) in 17 mL of demineralized water. After dry impregnation of 40 grams of support D’, the impregnated alumina is left to mature in a water-saturated atmosphere for 4 hours at room temperature, then dried at 120°C for 4 hours. The catalyst thus obtained is denoted D.

The final composition of elements of catalyst D, expressed in the form of oxides and relative to the weight of the dry catalyst is then as follows: MoO ₃ = 5.2 +/- 0.2% by weight, CoO = 1.1 +/- 0.1% by weight and P ₂ O ₅ = 0.7 +/- 0.1% by weight.

The Co/Mo and P/Mo molar ratios are 0.42 and 0.27, respectively.

The specific surface area of catalyst D is 189 ^m2 /g.

Example 5: Preparation of catalyst E

A support E’ identical to support C’ is provided.

Cobalt and molybdenum are then added. The impregnation solution is prepared by dissolving at room temperature ammonium heptamolybdate tetrahydrate (5.64 g, ≥99.5%, Sigma-Aldrich®), cobalt nitrate hexahydrate (5.36 g, ≥99.5%, Alfa Aesar®), in 28 mL of demineralized water. After dry impregnation of 40 grams of support E’, the impregnated alumina is left to mature in a water-saturated atmosphere for 4 hours at room temperature, then dried at 120°C for 4 hours, and finally calcined under an air flow rate of 1L/h/g at 450°C for 4 hours. The catalyst thus obtained is denoted E.

The final composition of elements of catalyst E, expressed in the form of oxides and relative to the weight of the dry catalyst is then as follows: MoO ₃ = 10.0 +/- 0.2% by weight and CoO = 3.0 +/- 0.1% by weight.

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

The specific surface area of catalyst E is 124 ^m2 /g.

Example 6: Implementation of catalysts in a gasoline desulfurization process

A gasoline from a catalytic cracking unit composed of 25% by weight of olefins and 600 ppmS in total sulfur is subjected to a 2-stage treatment:

- a first stage of selective hydrodesulfurization (HDS) in an adiabatic reactor using catalyst E. The operating conditions of the single-stage hydrodesulfurization stage of the gasoline feedstock are as follows: VVH = 3 h ^-1 , P = 2.0 MPa. A flow of pure hydrogen is added to the feedstock at the reactor inlet such that H ₂ /HC = 250 Nm ³ /m ³ . The effluent is sent directly to the reactor of the second stage;

- a second finishing hydrodesulfurization (FNS) stage in an adiabatic reactor using catalysts A to D. Only the effluent from the first stage is treated in this second stage. The operating conditions of the finishing stage are as follows: VVH = 3 h ^-1 , P = 2.0 MPa. The temperature at the reactor inlet is always set at 35°C higher than the temperature of the effluent at the outlet of the first selective hydrodesulfurization stage.

The inlet temperature into the reactor of the first stage of selective hydrodesulfurization is set to obtain an effluent containing 10 ppm wt S in total sulfur (i.e. more than 98% conversion into total sulfur)

Prior to use, the catalysts in the selective hydrodesulfurization and finishing reactors are 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 dimethyldisulfide in n-heptane.

The properties of the catalysts are presented in Table 1. The performances in gasoline desulfurization process are presented in Table 2.

The results illustrate that the process according to the invention for the hydrodesulfurization of gasoline makes it possible to reduce the temperature levels of treatment of the gasoline feedstock in the first HDS stage compared to the implementation of catalysts known from the prior art and therefore allows a reduction in the cycle duration, while maintaining high desulfurization rates and a limitation of olefin losses and therefore a maintenance of product specifications.

Catalyst A (compliant) B (compliant) C (non-compliant) D (non-compliant) % weight MoO ₃ 9.5 8.2 0 5.2 % CoO weight 1.8 1.5 0 1.1 % weight NiO 0 0 17.9 0 % weight P ₂ O ₅ 1.3 1.0 0 0.7 P/VIB
(mol/mol) 0.27 0.25 - 0.27 VIII/VIB
(mol/mol) 0.37 0.36 - 0.42 Specific surface area
(m ² /g) 31 75 114 189

Catalyst used in the finishing stage A (compliant) B (compliant) C (non-compliant) D (non-compliant) Olefins entry stage 1
(% weight) 25.0 25.0 25.0 25.0 Olefins output stage 2
(% weight) 16.7 16.2 16.5 15.3 Overall hydrogenation of olefins (%) 33.2 35.2 34.0 38.2 Input sulfur content (ppm wt S) 600 600 600 600 Output sulfur content (ppm wt S) 10 10 10 10 Reactor inlet temperature stage 1 (°C) 228 223 253 219 Reactor outlet temperature stage 1 (°C) 239 228 289 222 Average temperature step 1 (°C) 234 226 271 221 Reactor inlet temperature stage 2 (°C) 274 263 324 257 Reactor outlet temperature stage 2 (°C) 300 294 326 295

The “average temperature” corresponds to the Weight Average Bed Temperature (WABT) according to the consecrated Anglo-Saxon term, well known to those skilled in the art. The average temperature is advantageously determined according to the catalytic systems, equipment, and their configuration, used. The average temperature (or WABT) is calculated as follows:

with T _inlet : the temperature of the flow entering the reaction section and T _outlet : the temperature of the effluent leaving the reaction section. Unless otherwise indicated, the “average temperature” of a reaction section is given at start-of-cycle conditions.

Claims (10)

A process for treating gasoline containing sulfur compounds and olefins, the process comprising at least the following steps:
a) bringing into contact gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a metal from group VIB and a metal from group VIII at least partly in sulfurized form, and an oxide support, at a temperature between 200°C and 350°C, at a pressure between 0.2 MPa and 5 MPa, with an hourly volumetric flow rate between 1 h ^-1 and 20 h ^-1 and a ratio between the hydrogen flow rate expressed in normal m ³ per hour and the volumetric flow rate of the feedstock to be treated expressed in m ³ per hour under standard conditions between 10 Nm ³ /m ³ and 1000 Nm ³ /m ³ , to obtain a partially desulfurized effluent;
b) without separation of the H ₂ S formed in step a), the partially desulfurized effluent obtained at the end of step a) is brought directly into contact with a finishing hydrodesulfurization catalyst at a temperature of between 250°C and 400°C, at a pressure of between 0.2 MPa and 5 MPa, with an hourly volumetric flow rate of between 1 and 20 h ^-1 , to obtain a desulfurized effluent, said finishing hydrodesulfurization catalyst comprising an active phase comprising, preferably consisting of, at least one metal from group VIB and at least one metal from group VIII at least partly in sulfurized form, phosphorus, and a porous support based on alumina, the content of metal from group VIB, measured in oxide form, being between 1 and 20% by weight relative to the total weight of the catalyst, the content of metal from group VIII, measured in oxide form, being between 0.2 and 10% by weight relative to the total weight of the catalyst, and the phosphorus content, measured in its P ₂ O ₅ form, being between 0.1 and 5% by weight relative to the total weight of the catalyst, said catalyst having a surface area greater than or equal to 20 m²/g and less than 150 m ² /g. The process of claim 1, wherein the catalyst of step b) comprises a molar ratio between said group VIII metal of the active phase and said group VIB metal of the active phase is between 0.1 and 2.0 mol/mol. Process according to one of claims 1 or 2, in which the catalyst of step b) comprises a molar ratio between phosphorus and said metal of group VIB of the active phase is between 0.1 and 2.0 mol/mol. A process according to any preceding claim, wherein the catalyst of step b) comprises a specific surface area between 20 m ² /g and 110 m ² /g. A process according to any preceding claim, wherein the catalyst of step b) comprises cobalt as a group VIII metal and molybdenum as a group VIB metal. Process according to any one of the preceding claims, in which the catalyst of step b) comprises an active phase consisting of molybdenum, cobalt and phosphorus, and a porous support based on alumina, the cobalt content being between 0.5 and 5% by weight, measured in the oxide form CoO, relative to the total weight of the catalyst, the molybdenum content being between 3 and 12% by weight, measured in the oxide form MoO ₃ , relative to the total weight of the catalyst, a phosphorus content of between 0.3 and 3% by weight, measured in the oxide form P ₂ O ₅ , relative to the total weight of the catalyst, the molar ratio between cobalt and molybdenum being between 0.3 and 1.0 mol/mol, the molar ratio between phosphorus and molybdenum being between 0.2 and 0.5 mol/mol, the specific surface area of the catalyst being between 25 ^m2 /g and 90 ^m2 /g. Process according to any one of the preceding claims, in which the catalyst of step a) comprises an alumina support and an active phase comprising cobalt and molybdenum, 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% by weight, and a content by weight relative to the total weight of catalyst of molybdenum oxide, in MoO ₃ form, of between 1 and 20% by weight, with a cobalt/molybdenum molar ratio of between 0.1 and 0.8 mol/mol. A process according to any preceding claim, wherein the catalyst of step a) comprises a specific surface area of between 60 and 250 m²/g. A method according to any preceding claim, wherein the temperature of step b) is higher than the temperature of step a). A process according to any preceding claim, wherein the gasoline is a catalytic cracked gasoline.