WO2024115275A1

WO2024115275A1 - Process for finishing hydrodesulfurization of gasolines using a chain of catalysts

Info

Publication number: WO2024115275A1
Application number: PCT/EP2023/082853
Authority: WO
Inventors: Marie DEHLINGER; Damien Hudebine; Alexandre VONNER; Charlie BLONS; Antoine Fecant
Original assignee: IFP Energies Nouvelles
Priority date: 2022-11-30
Filing date: 2023-11-23
Publication date: 2024-06-06
Also published as: FR3142486A1

Abstract

Process for treating a gasoline containing sulfur compounds and olefins, comprising the following steps: a) in a first reaction section, contacting the gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a group VIB metal and a group VIII metal, and a support, to obtain a first partially desulfurized effluent; b) in a second reaction section, contacting the first partially desulfurized effluent and a first finishing hydrodesulfurization catalyst comprising an active phase of a group VIII metal, and a support, to obtain a second partially desulfurized effluent; c) in a third reaction section, contacting the second partially desulfurized effluent and a second finishing hydrodesulfurization catalyst comprising an active phase comprising a group VIB metal and a group VIII metal, and a support, to obtain a third desulfurized effluent.

Description

HYDRODESULFURIZATION PROCESS FOR FINISHING GASOLINES USING A SEQUENCE OF CATALYSTS

Field of the 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 gasoline with a low sulfur content. The invention applies particularly to the treatment of gasoline cuts containing olefins and sulfur, such as gasolines resulting from catalytic cracking, for which we seek to reduce the content of sulfur compounds, without hydrogenating the olefins and aromatics.

State of the art

The specifications for automobile fuels require a significant reduction in the sulfur content in these fuels, and in particular in gasoline. This reduction is intended to limit, in particular, the content of sulfur and nitrogen oxide in automobile exhaust gases. The specifications currently in force in Europe since 2009 for gasoline fuels set a maximum content of 10 ppm by weight (parts per million) of sulfur. Such specifications are also in force in other countries such as for example the United States and China where the same maximum sulfur content has been required since January 2017. To achieve these specifications, it is necessary to treat the gasolines with desulfurization 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 of atmospheric or vacuum distillation. empty of crude oil. The fraction of gasoline resulting from catalytic cracking, which represents on average 40% of gasoline bases, in fact contributes more than 90% to the sulfur content in gasolines. Consequently, the production of low sulfur gasolines requires a catalytic cracking gasoline desulfurization step. Other sources of gasoline that may contain sulfur also include coker gasoline, visbreaker gasoline or, to a lesser extent, gasoline from atmospheric distillation or steam cracking gasoline.

The elimination of sulfur in gasoline cuts consists of specifically treating these gasolines rich in sulfur by desulfurization processes in the presence of hydrogen. We then speak of hydrodesulfurization (HDS) processes. However, these essence cuts and more particularly gasolines resulting from catalytic cracking in a fluidized bed (or FCC for Fluid Catalytic Cracking according to Anglo-Saxon 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 number, diolefins (0.5 to 5% by weight) and aromatics. These unsaturated compounds are unstable and react during the hydrodesulfurization treatment. Diolefins form gums by polymerization during hydrodesulfurization treatments. This formation of gums leads to progressive deactivation of the hydrodesulfurization catalysts or progressive clogging of the reactor. Consequently, diolefins must be eliminated by hydrogenation before any treatment of these essences. Traditional treatment processes desulphurize gasolines in a non-selective manner by hydrogenating a large part of the mono-olefins, which causes a significant 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 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 the cracked gasolines must be desulfurized very deeply, part of the olefins present in the cracked gasolines are hydrogenated on the one hand and recombine with the F^S to form mercaptans of others share. This family of compounds, with 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 content of recombinant mercaptans can be carried out by catalytic hydrodesulphurization, but this leads to the hydrogenation of a significant part of the mono-olefins present in the gasoline, which then leads to a sharp reduction in the carbon dioxide index. octane in gasoline as well as overconsumption of hydrogen. It is also known that the loss of octane linked to the hydrogenation of mono-olefins during the hydrodesulfurization step is all the greater as the targeted sulfur content is low, i.e. that we seek to thoroughly eliminate the sulfur compounds present in the load.

It is thus possible to treat gasoline by a sequence of two reactors as described in document EP1077247, the first stage, also called the selective HDS stage, generally aims to carry out deep desulphurization of gasoline with minimal 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 H2S resulting from desulfurization with olefins: recombination mercaptans. The second step generally has the role of minimizing the quantity of recombinant mercaptans. The temperature is generally higher in the second stage in order to thermodynamically favor the elimination of mercaptans. In practice, an oven 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 the 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.

Nevertheless, 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 sequence of two specific catalysts of different nature and in a particular order in the finishing hydrodesulfurization section located downstream of the selective hydrodesulfurization section (HDS) presents a synergistic effect in terms of selectivity by minimizing the saturation of the olefins which leads to a loss in octane number. Indeed, the choice of the right active phase and a suitable support makes it possible to observe a synergy between a first catalyst in the finishing section consisting of an active phase based on a group VIII element, allowing the elimination of recombinant mercaptans while preserving the olefins and a second catalyst in the finishing section consisting of an active phase of a group VIII element and a group VIB element, allowing the elimination of more refractory sulfur compounds . On the other hand, the introduction of a second catalyst in the finishing section consisting of an active phase of an element from group VIII and an element from group VIB, makes it possible to reduce the average treatment temperature in the HDS section and thus increase the overall cycle time of the process. Objects of the invention

The aim of the pre-sale invention is to implement a process for producing gasoline with a low sulfur content, making it possible to recover the entirety of a gasoline cut containing sulfur, preferably a catalytic cracked gasoline cut, and to reduce the sulfur contents in said gasoline cut at very low levels, without reduction in gasoline yield while minimizing the reduction in the octane number due to the hydrogenation of olefins.

Thus, the subject of the present invention is a process for treating a gasoline containing sulfur compounds and olefins, the process comprising at least the following steps: a) the gasoline is brought into contact in a first reaction section, hydrogen and a hydrodesulphurization catalyst comprising an active phase comprising a Group VI B metal and a Group VIII metal at least partly in sulphurized 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 volume 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 feed rate to be treated expressed in m ³ per hour at standard conditions of between 10 Nm ³ /m ³ and 1000 Nm ³ /m ³ , to obtain a first partially desulfurized effluent; b) without separation of the H2S formed in step a), the first partially desulfurized effluent obtained at the end of step a) is directly brought into contact in a second reaction section, and a first hydrodesulfurization catalyst finishing comprising, an active phase consisting of a Group VIII metal at least partly in sulfide form, and an oxide support, at a temperature between 250°C and 400°C, at a pressure between 0, 2 MPa and 5 MPa, with an hourly volume velocity of between 1 h ^-1 and 40 h' ¹ , to obtain a second partially desulfurized effluent; c) without separation of the H ₂ S formed in step b), the second partially desulfurized effluent obtained at the end of step b) is directly brought into contact in a third reaction section, and a second catalyst d finishing hydrodesulphurization comprising an active phase comprising, preferably consisting of, at least one metal from group VI B and at least one metal from group VIII at least partly in sulphided form, and an oxide support, at a temperature comprised between 250°C and 400°C, at a pressure of between 0.2 MPa and 5 MPa, with an hourly volume velocity of between 1 ^h'1 and 40 ^h'1 , to obtain a third desulfurized effluent.

According to one or more embodiments, said second reaction section containing the first finishing hydrodesulfurization catalyst occupies a volume V1, and said third finishing hydrodesulfurization reaction section containing the second finishing hydrodesulfurization catalyst occupies a volume V2, the distribution of volumes V1/V2 being between 90%vol/10%vol and 10%vol/90%vol, preferably between 90%vol/10%vol and 60%vol/40%vol, respectively of said second and third finishing hydrodesulfurization reaction section.

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

According to one or more embodiments, the catalyst of step a) and/or step c) comprises alumina and an active phase comprising cobalt and molybdenum, said catalyst containing a content by weight relative to to the total weight of cobalt oxide catalyst, in CoO form, of between 0.1 and 10% and a content by weight relative to the total weight of molybdenum oxide catalyst, in MoOs form, of between 1 and 20 %, 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) and/or step c) further comprises phosphorus, said catalyst containing a content by weight relative to the total weight of phosphorus oxide catalyst in P2O5 form between 0.3 and 10% by weight.

According to one or more embodiments, the catalyst of step a) and/or step c) comprises a specific surface area between 60 and 250 m ² /g.

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

According to one or more embodiments, the catalyst of step b) comprises a Group VIII metal content of between 5 and 65% by weight of Group VIII metal oxide relative to the total weight of the catalyst.

According to one or more embodiments, the catalyst of step b) comprises an alumina support and an active phase consisting of nickel, said catalyst containing a content by weight relative to the total weight of catalyst of nickel oxide, in NiO form, between 5 and 65% by weight. According to one or more embodiments, the catalyst of step b) comprises a specific surface area between 60 and 250 m ² /g.

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

According to one or more embodiments, steps b) and c) are carried out in a single reactor.

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

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

Detailed description of the invention

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 IIIPAC classification.

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

In the following discussion of the invention, by total pore volume of the oxide support or catalyst is meant the volume measured by intrusion with a mercury porosimeter according to standard 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 equal to 140° following the recommendations of the work “Engineering techniques, analysis and characterization treatise”, p.1050-5, written by Jean Charpin and Bernard Rasneur.

In order to obtain better precision, the value of the total pore volume in ml/g or cm ³ /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 in cm ³ /g measured on the sample minus the value of the mercury volume in ml/g or in cm ³ /g measured on the same sample for a pressure corresponding to 30 psi (approximately 0.2 MPa).

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

The contents of group VIB metals, group VIII elements 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. Loss on ignition is due to loss of moisture. It is determined according to ASTM D7348.

Load

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 (coking according to Anglo-Saxon terminology), visbreaking (visbreaking according to Anglo-Saxon terminology), steam cracking (according to Anglo-Saxon terminology) or catalytic cracking (FCC, Fluid Catalytic Cracking according to Anglo-Saxon terminology). This gasoline may possibly be composed of a significant fraction of gasoline coming from other production processes such as atmospheric distillation (gasoline resulting from direct distillation (or straight run gasoline according to Anglo-Saxon terminology)) or from processes conversion (coking or steam cracking gasoline). Said feed preferably consists of a gasoline cut from a catalytic cracking unit.

The feed 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 at 220°C. The method according to the invention can also treat fillers having end points lower than those mentioned above, such as for example a C5-180°C cut.

The sulfur content of gasoline cuts produced by catalytic cracking (FCC) depends on the sulfur content of the feed treated by the FCC, the presence or absence of pretreatment of the FCC feed, as well as the end point of the cut. Generally, the sulfur contents of an entire gasoline cut, particularly those coming from FCC, are greater than 100 ppm by weight and most of the time greater than 500 ppm by weight. For species with end points greater than 200°C, sulfur contents are often greater than 1000 ppm by weight, they can even in certain cases reach values of the order of 4000 to 5000 ppm by weight. The feed treated by the process according to the invention may be a feed 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% weight of sulfur including generally less than 300 ppm of mercaptans.

Step aO) Selective hydrogenation (optional)

Depending on the type of gasoline to be treated, it may be advantageous to first 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 weighting reaction. part of the light mercaptans (RSH) present in the thioethers charge, by reaction with olefins.

To this end, the gasoline to be treated is sent to a catalytic selective hydrogenation reactor containing at least one fixed or mobile bed of catalyst for the selective hydrogenation of diolefins and the weighting of light mercaptans. The reaction for selective hydrogenation of diolefins and weighting of light mercaptans is preferably carried out on a sulfide catalyst comprising at least one element from group VIII and optionally at least one element from group VIB and an oxide support. The group VIII element is preferably chosen from nickel and cobalt and in particular nickel. The group VIB element, 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. Preferably, alumina is 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 NiO form, of between 1 and 12%, and molybdenum at a content by weight of oxide of molybdenum, in MoOs form, 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%.

During the optional selective hydrogenation step, the gasoline is brought into contact with the catalyst at a temperature between 50 and 250°C, and preferably between 80 and 220°C, and even more preferably between 90°C. and 200°C, with an hourly volume velocity (WH) of between 0.5 h ^-1 and 20 h' ¹ , the unit of the hourly volume velocity being the volume flow rate of charge at 15°C per volume of catalytic bed (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 m ³ per hour and the volume flow rate of feed to be treated expressed in m ³ per hour at standard conditions (15°C, 0 .1 MPa) between 2 and 100 Nm ³ /m ³ , preferably between 3 and 30 Nm ³ /m ³ .

After selective hydrogenation, the diolefin content, determined via the maleic anhydride index (MAV or “Maleic Anhydride Value” according to Anglo-Saxon terminology), according to the UOP 326 method, is generally reduced to less than 6 mg of maleic anhydride/g, or even less than 4 mg AM/g and more preferably less than 2 mg AM/g. In some cases, less than 1 mg AM/g can be obtained.

The selectively hydrogenated gasoline can then be distilled into at least two cuts, a light cut and a heavy cut and possibly an intermediate cut. In the case of splitting into two cuts, the heavy cut is treated according to the process of the invention. In the case of splitting 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 stages of hydrogenation of diolefins and fractionation in 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 (H DS)

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 m ³ per hour, per m ³ of catalytic bed (also called hourly volume velocity - WH) 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 load 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 . By normal m ³ we mean the volume of 1 m ³ of gas at 0°C and 0.1 MPa. The hydrogen required for this step can be fresh hydrogen or recycled hydrogen, preferably free of 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 resulting 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 rate of hydrogenation of the olefins is preferably less than 50%, more preferably less than 40% during this step.

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

The group VI B metal present in the active phase of the catalyst is preferably chosen from molybdenum and tungsten.

The Group VIII metal present in the active phase of the catalyst is preferably chosen from cobalt, nickel and the mixture of these two elements.

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

The content of Group VIII metal is preferably between 0.1 and 10% by weight of oxide of the Group VIII metal 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 content of metal from group VI B is preferably between 1 and 20% by weight of oxide of the metal from group VI B 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 oxide of the metal from Group VI B relative to the total weight of the catalyst. When the metal is molybdenum or tungsten, the metal content is expressed as MoOs or WO3. Preferably, the molar ratio of Group VIII metal to Group VI B 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 also have a phosphorus content generally between 0.3 and 10% by weight of P2O5 relative to the total weight of catalyst, preferably between 0.3 and 5% by weight, very preferably between 0 .5 and 3% weight.

Furthermore, when phosphorus is present, the phosphorus/(group VI B 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 m ² /g, preferably between 60 and 200 m ² /g, and even more preferably between 65 and 180 m ² /g , and even more preferably between 70 and 130 m ² /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 chosen from the group consisting of: alumina, silica, silica-alumina or even titanium or magnesium oxides used alone or in mixture with the alumina or silica-alumina. It is preferably chosen from the group consisting of silica, alumina, and silica-alumina. Very preferably, the oxide support consists essentially of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80%. weight, or even at least 90% by weight of alumina relative to the total weight of said oxide support. It preferably consists 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 and optionally phosphorus, said catalyst containing a content by weight per relative to the total weight of cobalt oxide catalyst, in CoO form, 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 molybdenum oxide catalyst, in MoOs form, of between 1 and 20% by weight, preferably between 2 and 18% by weight, and so very preferred between 3 and 16% 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 m ² /g, preferably between 60 and 200 m ² /g, and even more preferably between 65 and 180 m ² /g, and even more preferably between 70 and 130 m ² /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 for the hydrodesulfurization catalyst can be in the form of balls, extrudates of any geometry, platelets, pellets, compressed cylinders, crushed solids or any other format. Preferably, the support is in the form of balls of 0.5 to 6 mm in diameter or in the form of cylindrical, trilobed or quadrilobed extrudates of 0.8 to 3 mm in circumscribed diameter. More preferably, the support is in the form of balls.

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) First finishing hydrodesulfurization step (FNS1)

During hydrodesulfurization step a), a large part of the sulfur compounds are transformed into H2S. The remaining sulfur compounds are essentially refractory sulfur compounds and the recombination mercaptans resulting from the addition of the H2S formed in step a) to the olefins present in the feed.

Step b) of the process according to the invention consists of transforming at least part of the recombinant mercaptans contained in the first effluent from step a) into olefins and H2S as well as at least part of the sulfur compounds contained in the first effluent from step a) such as thiophene compounds, into compounds saturated for example with thiophanes (or thiacyclopentanes) or mercaptans, then at least partially hydrogenolyzing these sulfur compounds to form H2S.

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 equilibrium thermodynamics. 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 of the second reaction section is generally such that the ratio between the volume flow rate of gasoline to be treated expressed in m ³ per hour at standard conditions (15°C, 0.1 MPa), per m ³ of catalytic bed (also called hourly volumetric speed - WH) is between 1 and 40 h ^-1 and preferably between 2 and 20 h ^-1 .

The first hydrodesulphurization catalyst of step b) comprises an active phase consisting of a Group VIII metal, and an oxide support, said active phase being at least partly sulphurized.

The Group VIII metal is preferably nickel. When the Group VIII metal is nickel, the phase diagram of nickel sulfide shows a large number of sulfur-rich and nickel-rich phases at low temperatures. Different phases and stoichiometries of nickel sulfide are therefore possible, ranging from nickel-rich compounds such as Ni ₃ S2, NieSs, Ni ₇ Se, NigSs and NiS to sulfur-rich compounds like Ni ₃ S4 and NiS2. Note that NiS is also known to exist in two main phases, namely hexagonal a-NiS, stable at high temperatures, and rhombohedral p-NiS stable at low temperatures. The existence of these numerous phases makes the synthesis of nickel sulfide in the form of a single phase complex, the products therefore often being mixtures of two or more phases.

The content of Group VIII metal is preferably between 5 and 65% by weight of oxide of the Group VIII metal relative to the total weight of the catalyst, more preferably between 8 and 55% by weight, very preferably between 12 and 40% by weight of Group VIII metal oxide relative to the total weight of the catalyst. When the metal is nickel, the metal content is expressed as NiO.

Preferably, the catalyst of step b) is characterized by a specific surface area of between 60 and 250 m ² /g, preferably between 70 and 200 m ² /g. The total pore volume of the catalyst of step b) 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 first finishing hydrodesulfurization catalyst is typically a porous solid chosen from the group consisting of: alumina, silica, silica-alumina or even titanium or magnesium oxides used alone or in a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, alumina, and silica-alumina. Very preferably, the oxide support consists essentially of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80%. weight, or even at least 90% by weight of alumina relative to the total weight of said oxide support. It preferably consists solely of alumina.

Preferably, the support of the first finishing hydrodesulfurization catalyst comprises a specific surface area of between 60 and 250 m ² /g, preferably between 70 and 200 m ² /g.

The total pore volume of the support of the first finishing hydrodesulfurization catalyst 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 for the first finishing hydrodesulfurization catalyst can be in the form of balls, extrudates of any geometry, platelets, pellets, compressed cylinders, crushed solids or any other format. Preferably, the support is in the form of balls of 0.5 to 6 mm in diameter or in the form of cylindrical, trilobed or quadrilobed extrudates of 0.8 to 3 mm in circumscribed diameter. More preferably, the support is in the form of balls.

The second partially desulfurized effluent obtained at the end of step b) is then sent directly and without separation to step c) of the process according to the invention.

Step c) Second stage of finishing hydrodesulfurization (FNS2)

The sulfur compounds remaining at the end of step b) are essentially refractory sulfur compounds. Step c) of the process according to the invention essentially consists of transforming at least part of the refractory sulfur compounds contained in the effluent resulting from step b) such as thiophenic compounds, into saturated compounds for example into thiophanes (or thiacyclopentanes) or mercaptans, then at least partially hydrogenolyze these sulfur compounds to form H2S. 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 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 m ³ per hour at standard conditions (15°C, 0.1 MPa), per m ³ of bed catalytic (also called hourly volume velocity) is between 1 and 40 h ^-1 and preferably between 2 and 20 h' ¹ .

In the process according to the invention the rate of hydrogenation of the olefins of steps b) and c) is preferably less than 30% during these steps.

The total desulfurization rate of steps b) and c), which depends on the sulfur content of the feed to be treated, is generally greater than 50% and preferably greater than 70% so that the product resulting from step c ) 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 second finishing hydrodesulfurization catalyst of step c) comprises an active phase comprising, preferably consisting of, at least one metal from Group VI B and at least one metal from Group VIII at least partly in sulfide form. optionally phosphorus, and an oxide support, as described below.

The group VI B metal is preferably chosen from molybdenum and tungsten. The group VIII metal is preferably chosen from cobalt, nickel and the mixture of these two elements. The active phase of the catalyst is preferably chosen from the group formed by the combination of the elements nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum and very preferably the active phase consists of cobalt and molybdenum.

The content of Group VIII metal is preferably between 0.1 and 10% by weight of oxide of the Group VIII metal 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. When the metal is cobalt or nickel, the metal content is expressed as CoO or NiO. The metal content of group VIB is preferably between 1 and 20% by weight of oxide of the metal of group VIB relative to the total weight of the catalyst, more preferably between 2 and 18% by weight, and very preferably between 3 and 16% weight. When the metal is molybdenum or tungsten, the metal content is expressed as MoOs or WO3.

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 also have a phosphorus content generally between 0.3 and 10% by weight of P2O5 relative to the total weight of catalyst, preferably between 0.3 and 5% by weight, very preferably between 0 .5 and 3% 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 c) comprises a specific surface area of between 60 and 250 m ² /g, preferably between 60 and 200 m ² /g, and even more preferably between 65 and 180 m ² /g , and even more preferably between 70 and 130 m ² /g.

The total pore volume of the catalyst of step c) 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.

In a preferred embodiment, the catalyst of step c) 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 cobalt oxide catalyst, in CoO form, 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 molybdenum oxide catalyst, in the form MoOs, 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 preferably between 0.2 and 0.6 mol/mol.

Preferably, the support of the second finishing hydrodesulfurization catalyst comprises a specific surface area of between 60 and 250 m ² /g, preferably between 60 and 200 m ² /g, and even more preferably between 65 and 180 m ² /g, and even more preferably between 70 and 130 m ² /g.

The total pore volume of the support of the second finishing hydrodesulfurization catalyst 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 for the second finishing hydrodesulfurization catalyst can be in the form of balls, extrudates of any geometry, platelets, pellets, compressed cylinders, crushed solids or any other format. Preferably, the support is in the form of balls of 0.5 to 6 mm in diameter or in the form of cylindrical, trilobed or quadrilobed extrudates of 0.8 to 3 mm in circumscribed diameter. More preferably, the support is in the form of balls.

In a preferred variant, the catalyst used in step c) is the same as that used in step a).

Preferably step c) is carried out in the same reactor as step b).

Implementation of steps b) and c) of the process

Steps b) and c) of the process according to the invention can be carried out in one, two or more reactors.

When steps b) and c) of the process according to the invention are carried out in two different reactors, step b) can be carried out in a first finishing hydrodesulfurization reactor containing the second reaction section through which the partially desulfurized effluent passes. from step a), then step c) can be carried out in the second finishing hydrodesulfurization reactor containing the third reaction section, placed downstream of said first reactor. When steps b) and c) of the process according to the invention are carried out in a single reactor, step b) is carried out in a first zone containing the second reaction section, and step c) is carried out in a second zone containing the third reaction section downstream of the first zone.

According to one or more embodiments, said second reaction section containing the first finishing hydrodesulfurization catalyst occupies a volume V1, and said third finishing hydrodesulfurization reaction section containing the second finishing hydrodesulfurization catalyst occupies a volume V2, the distribution of volumes V1/V2 being between 90%vol/10%vol and 10%vol/90%vol respectively of said second and third finishing hydrodesulphurization reaction section, preferably between 90%vol/10%vol and 60%vol/40%vol, respectively of said second and third finishing hydrodesulfurization reaction section.

Step d): Separation of H ₂ S (optional)

Separation step d) is carried out in order to separate the excess hydrogen as well as the H ₂ S formed during steps a), b) and c). Any method known to those skilled in the art can be considered.

According to a first embodiment, after steps a), b) and c), the third desulfurized effluent from step c) 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 flask. 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 made up of residual H ₂ S and hydrocarbon compounds having a boiling temperature lower than or equal to that of butane and a bottom cut free of H ₂ S, called stabilized gasoline, containing 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 hydrogen and H ₂ S is sent to a purification section. 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 dissolved in the liquid fraction. as well as the residual dissolved 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 d) is carried out in a stabilization column or debutanizer. In fact, a stabilization column makes it possible to separate H2S more efficiently than a stripping section.

Step d) is preferably implemented so that the sulfur in the form of H2S remaining in the desulphurized 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 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 elements of groups VIII and possibly VIB and 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 quantity of precursors of the desired elements is introduced in the form of salts soluble in the chosen solvent, for example demineralized water, 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 favoring 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 an organic agent mixed with the precursors of the elements of group VIII, group VI and phosphorus.

By way of example, among the sources of molybdenum, one can use oxides and hydroxides, molybdic acids and their salts, in particular ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PM012O40), and their salts, and optionally silicomolybdic acid (hLSiMo^C o) and its salts. The sources of molybdenum can also be any heteropolycompound of the Keggin type, lacunar Keggin, substituted Keggin, Dawson, Anderson, Strandberg, for example. Molybdenum trioxide and heteropolycompounds of the Keggin, lacunar Keggin, substituted Keggin and Strandberg type are preferably used.

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

The nickel precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example. Nickel hydroxide and nickel hydroxycarbonate are preferably used. The tungsten precursors which can be used are also well known to those skilled in the art. For example, among the tungsten sources, one can use oxides and hydroxides, tungstic acids and their salts, in particular ammonium salts such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and their salts. salts, and possibly silicotungstic acid (H ₄ SiWi ₂ 04o) and its salts. The sources of tungsten can also be any heteropolycompound of the Keggin type, lacunar Keggin, substituted Keggin, Dawson, for example. Ammonium oxides and salts such as ammonium metatungstate or heteropolyanions of the Keggin, lacunar Keggin or substituted Keggin type are preferably used.

Phosphorus can advantageously be introduced alone or mixed with at least one of the elements of group VI B and VIII. The phosphorus is preferably introduced as a mixture with the precursors of the elements of group VI B and of group VIII by dry impregnation of said porous support using a solution containing the precursors of the elements and the phosphorus precursor. The preferred source of phosphorus is orthophosphoric acid H3PO4, but its salts and esters such as ammonium phosphates or mixtures thereof are also suitable. Phosphorus can also be introduced at the same time as the element(s) of group VI B in the form, for example, of Keggin, lacunar Keggin, substituted Keggin or Strandberg type heteropolyanions.

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 step, the catalyst precursor obtained can undergo 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 simple drying of the catalyst can also be carried out.

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

In the case of an oxidizing treatment, also called calcination, this is generally carried out under air or diluted oxygen, and the treatment temperature is generally between 200°C and 550°C, preferably between 300°C. C and 500°C, and advantageously for a period typically between 0.5 hour to 24 hours, preferably for a period of 0.5 hour to 12 hours, and even more preferably for a period of 0.5 hour 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 sulfo-reducing atmosphere in the presence of hydrogen and hydrogen sulfide. Hydrogen sulfide can be used directly or generated by a sulfide agent (such as dimethyldisulfide).

Description of catalyst sulfurization

Before being brought into contact with the feedstock to be treated in a gasoline hydrodesulfurization process, the catalysts used in steps, optionally aO), a), b) and c) of the process according to the invention generally undergo a step of sulfurization. Sulfurization is preferably carried out in a sulforeductive medium, that is to say in the presence of H ₂ S and hydrogen, in order to transform the metal oxides into sulphides such as, for example, M0S2, CogSs or Ni ₃ S2. 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 dimethyldisulfide (DMDS) are H ₂ S precursors commonly used to sulfurize the catalysts of steps, optionally aO), a), b) and c). Sulfur can also come from the filler. The temperature is adjusted so that the H ₂ S 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°C and 600°C, and more preferably between 300°C and 500°C. .

The sulfurization rate of the metals constituting the catalysts of the steps, optionally aO), a), b) or c) is at least equal to 60%, preferably at least equal to 70%. The sulfur content in the sulfurized catalyst of steps, optionally aO), a), b) or c) is measured by elemental analysis according to ASTM D5373. A metal is considered to be sulphurized when the overall sulphurization rate defined by the molar ratio between the sulfur (S) present on the catalyst and said metal is at least equal to 60% of the theoretical molar ratio corresponding to the total sulphurization of the(s) metal(s) considered. The overall sulfurization rate is defined by the following equation:

(S/metal)catalyst — 0.6 X (S/metal)theoretical in which :

(S/metal)catai _y seur is the molar ratio between the sulfur (S) and the metal present on the catalyst (S/metal)theoretical is the molar ratio between the sulfur and the metal corresponding to the total sulfidation 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 used in step, optionally aO), a) or in step c) comprises several metals, the molar ratio between the sulfur 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 sulfidation of each metal to sulphide, the calculation being carried out in proportion to the relative mole fractions of each metal.

The examples below illustrate the invention without limiting its scope.

Examples

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

- sulfur content according to the 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 chromatography analysis according to ASTM D6733 method.

Example 1: Preparation of catalyst A

A support A' composed of alumina is provided in the form of beads with a particle size 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.

We then add cobalt and molybdenum. The impregnation solution is prepared by dissolving at room temperature ammonium heptamolybdate tetrahydrate (5.64 g, >99.5%, Sigma-Aldrich®) and cobalt nitrate hexahydrate (5.36 g, >99.5%, Sigma-Aldrich®) .5%, Alfa Aesar®) in 28 mL of demineralized water. After dry impregnation of 40 grams of support A', the impregnated alumina is left to mature in an atmosphere saturated with water for 4 hours at room temperature, then is dried at 120°C for 4 hours, and finally calcined at a flow rate of air of 1 L/h/g at 450°C for 4 hours. The catalyst thus obtained is denoted A. The final element composition of catalyst A, expressed in the form of oxides, and related to the weight of the dry catalyst is then as follows: MoOs = 10.0 +/- 0.2% by weight and CoO = 3.0 +/- 0.1% weight.

The Co/Mo molar ratio is 0.60 mol/mol.

The specific surface area of catalyst A is 124 m ² /g.

Example 2: Preparation of catalyst B

We provide a support B’ identical to the support A’.

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

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

The specific surface area of catalyst B is 114 m ² /g.

Example 3: Implementation of catalysts in a gasoline desulfurization process Example 3 aims to show the benefit of the gasoline desulfurization process by using a sequence of steps and specific catalysts for each step. Gasoline from a catalytic cracking unit composed of 25% by weight of olefins and 600 ppmS of total sulfur is subjected to treatment in several stages:

- a selective hydrodesulfurization step (HDS) in an adiabatic reactor using catalyst A. The operating conditions of the one-step hydrodesulfurization step of the gasoline feed are as follows: WH = 3 h' ¹ , P = 2.0 MPa. A stream of pure hydrogen is added to the reactor inlet feed such that H2/HC = 250 Nm ³ /m ³ . The effluent is sent directly to the second stage reactor;

- a first finishing hydrodesulfurization step (FNS1) in an adiabatic reactor using catalysts A or B. Only the effluent from the first step is treated in this second step. The pressure of the first stage of finishing hydrodesulfurization is set at 2.0 MPa. The reactor inlet temperature is always set at 35°C higher than the temperature of the first effluent leaving the selective hydrodesulfurization stage;

- optionally, a second finishing hydrodesulfurization step (FNS2) in an adiabatic reactor using catalysts A or B. Only the effluent from the step previous is treated in this second finishing hydrodesulfurization step. The pressure of the second stage of finishing hydrodesulfurization is set at 2.0 MPa. The reactor inlet temperature is equal to the temperature of the effluent leaving the previous step. The entry temperature into the reactor of the selective hydrodesulfurization step is set to obtain an effluent containing 10 ppm by weight S in total sulfur (i.e. more than 98% conversion to total sulfur). Prior to its use, the catalysts contained in the selective hydrodesulphurization and finishing reactors are sulphurized 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 as dimethyldisulfide (DMDS) in n-heptane.

The results illustrate that the gasoline hydrodesulfurization process according to the invention makes it possible to obtain the best performance compared to implementations known from the prior art, since the implementation according to the invention makes it possible to increase the olefin content at the process outlet while minimizing an increase in the average temperature in the H DS section, which makes it possible to increase the life of the catalysts. The performance in the gasoline desulfurization process is presented in Table 1.

Table 1

The “average temperature” of the HDS, FNS1 or FNS2 stage corresponds to the Weight Average Bed Temperature (WABT) according to the established Anglo-Saxon term, well known to those skilled in the art. The average temperature is advantageously determined as a function of the catalytic systems, equipment and configuration thereof used. The average temperature (or WABT) is calculated as follows:

WABT = (Tinlet + Tj _ftt jj _iS ) 2 with Tinlet: the temperature of the flow entering the reaction section and Toutlet ■ the temperature of the effluent leaving the reaction section. Unless otherwise stated, the “average temperature” of a reaction section is given at cycle start conditions.

Claims

1. Process for treating a gasoline containing sulfur compounds and olefins, the process comprising at least the following steps: a) the gasoline, hydrogen and a catalyst are brought into contact in a first reaction section. hydrodesulphurization comprising an active phase comprising a Group VI B metal and a Group VIII metal at least partly in sulphide form, and an oxide support, at a temperature between 200°C and 350°C, at a pressure between between 0.2 MPa and 5 MPa, with an hourly volume 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 volume flow rate of the load to be treated expressed in m ³ per hour at standard conditions between 10 Nm ³ /m ³ and 1000 Nm ³ /m ³ , to obtain a first partially desulfurized effluent; b) without separation of the H2S formed in step a), the first partially desulfurized effluent obtained at the end of step a) is directly brought into contact in a second reaction section, and a first hydrodesulfurization catalyst finishing comprising an active phase consisting of a Group VIII metal at least partly in sulfide form, and an oxide support, at a temperature between 250°C and 400°C, at a pressure between 0.2 MPa and 5 MPa, with an hourly volume velocity of between 1 ^h'1 and 40 ^h'1 , to obtain a second partially desulfurized effluent; c) without separation of the H2S formed in step b), the second partially desulfurized effluent obtained at the end of step b) is directly brought into contact in a third reaction section, and a second hydrodesulfurization catalyst finishing comprising an active phase comprising at least one metal from group VIB and at least one metal from group VIII at least partly in sulfide form, and an oxide support, at a temperature between 250°C and 400°C, at a pressure of between 0.2 MPa and 5 MPa, with an hourly volume velocity of between 1 h ^-1 and 40 h' ¹ , to obtain a third desulfurized effluent.

2. Method according to claim 1, wherein said second reaction section containing the first finishing hydrodesulfurization catalyst occupies a volume V1, and said third finishing hydrodesulfurization reaction section containing the second finishing hydrodesulfurization catalyst occupies a volume V2, the distribution of volumes V1/V2 being between 90%vol/10%vol and 10%vol/90%vol respectively of said second and third reaction section.

3. Method according to one of claims 1 or 2, in which the catalyst of step a) and/or step c) comprises a group VIII metal content of between 0.1 and 10% by weight d oxide of the metal from Group VIII relative to the total weight of the catalyst, and a metal content from Group VI B of between 1 and 20% by weight of oxide of the metal from Group VI B relative to the total weight of the catalyst.

4. Process according to any one of the preceding claims, in which the catalyst of step a) and/or step c) comprises alumina and an active phase comprising cobalt and molybdenum, said catalyst containing a content by weight relative to the total weight of cobalt oxide catalyst, in CoO form, of between 0.1 and 10% and a content by weight relative to the total weight of molybdenum oxide catalyst, in MoOs form , between 1 and 20%, a cobalt/molybdenum molar ratio of between 0.1 and 0.8 mol/mol.

5. Method according to any one of the preceding claims, in which the catalyst of step a) and/or step c) further comprises phosphorus, said catalyst containing a content by weight relative to the total weight of phosphorus oxide catalyst in P2O5 form between 0.3 and 10% by weight.

6. Process according to any one of the preceding claims, in which the catalyst of step a) and/or step c) comprises a specific surface area between 60 and 250 m ² /g.

7. Process according to any one of the preceding claims, in which the catalysts of steps a) and c) are identical.

8. Process according to any one of the preceding claims, in which the catalyst of step b) comprises a Group VIII metal content of between 5 and 65% by weight of Group VIII metal oxide relative to the total weight. of the catalyst.

9. Method according to any one of the preceding claims, in which the catalyst of step b) comprises an alumina support and an active phase consisting of nickel, said catalyst containing a content by weight relative to the total weight of catalyst nickel oxide, in NiO form, between 5 and 65% by weight.

10. Process according to any one of the preceding claims, in which the catalyst of step b) comprises a specific surface area between 60 and 250 m ² /g.

11. Method according to any one of the preceding claims, in which before step a), the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst to selectively hydrogenate the diolefins contained in said gasoline into olefins.

12. Method according to any one of the preceding claims, in which steps b) and c) are carried out in a single reactor.

13. Method according to any one of the preceding claims, in which the temperature of steps b) and c) is higher than the temperature of step a).

14. Process according to any one of the preceding claims, in which the gasoline is a catalytic cracked gasoline.