Hydrodesulfurization process for gasoline finishing using a series of catalysts

Abstract

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

Description

Title of the invention: Hydrodesulfurization process for finishing gasoline using a chain of catalysts. Field of the invention

[0001] The present invention relates to the field of hydrotreating gasoline fractions, particularly gasoline fractions from fluidized bed catalytic cracking units. More specifically, the present invention concerns the use of catalysts in a low-sulfur gasoline production process. The invention is particularly applicable to the treatment of gasoline fractions containing olefins and sulfur, such as gasoline from catalytic cracking, for which the aim is to reduce the sulfur compound content without hydrogenating the olefins and aromatics. State of the art

[0002] Specifications for automotive fuels require a significant reduction in the sulfur content of these fuels, particularly gasoline. This reduction is intended to limit, in particular, the sulfur and nitrogen oxide content in vehicle exhaust gases. The specifications currently in force in Europe since 2009 for gasoline fuels set a maximum sulfur content of 10 ppm by weight (parts per million). Such specifications are also in force in other countries, such as the United States and China, where the same maximum sulfur content has been required since January 2017. To meet these specifications, it is necessary to treat gasoline using desulfurization processes.

[0003] The main sources of sulfur in gasoline bases are so-called cracked gasolines, and primarily, the gasoline fraction obtained from a catalytic cracking process of a vacuum distillate or a residue from the atmospheric or vacuum distillation of crude oil. The gasoline fraction from catalytic cracking, which represents on average 40% of gasoline bases, contributes more than 90% of the sulfur content in gasolines. Consequently, the production of low-sulfur gasolines requires a desulfurization step for the catalytic cracked gasolines. Other sources of gasoline that may contain sulfur include coker gasoline, visbreaker gasoline, and, to a lesser extent, gasolines from atmospheric distillation or steam cracking gasoline.

[0004] Sulfur removal from gasoline fractions involves specifically treating these sulfur-rich gasolines with desulfurization processes in the presence of hydrogen. These are known as hydrodesulfurization (HDS) processes. However, these gasoline fractions, and more specifically gasolines from catalytic bed cracking, are subject to these processes. Fluidized fuels (or FCC for Fluid Catalytic Cracking, according to Anglo-Saxon terminology) contain a significant proportion of unsaturated compounds in the form of monoolefins (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 process. Diolefins form gums through polymerization during hydrodesulfurization treatments. This gum formation leads to the progressive deactivation of the hydrodesulfurization catalysts or the gradual clogging of the reactor. Consequently, the diolefins must be removed by hydrogenation before any further treatment of these fuels. Traditional treatment processes desulfurize fuels non-selectively by hydrogenating a large portion of the monoolefins, resulting in a significant loss of octane rating and high hydrogen consumption.The most recent hydrodesulfurization processes make it possible to desulfurize cracked gasoline rich in monoolefins, while limiting the hydrogenation of the monoolefins and consequently the loss of octane. Such processes are described, for example, in documents EP-A-1077247 and EP-A-1174485.

[0005] However, when cracked gasoline undergoes very deep desulfurization, some of the olefins present in the cracked gasoline are hydrogenated and recombine with H₂S to form mercaptans. This family of compounds, with the chemical formula R-SH where R is an alkyl group, are generally called recombinant mercaptans and typically represent between 20% and 80% by weight of the residual sulfur in desulfurized gasoline. Reducing the content of recombinant mercaptans can be achieved by catalytic hydrodesulfurization, but this results in the hydrogenation of a significant portion of the monoolefins present in the gasoline, leading to a sharp decrease in the gasoline's octane rating and increased hydrogen consumption.It is also known that the octane loss linked to the hydrogenation of mono-olefins during the hydrodesulfurization step is greater the lower the target sulfur content, i.e., the aim is to thoroughly eliminate the sulfur compounds present in the feedstock.

[0006] 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 the 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 the H2S produced by desulfurization with the olefins: the recombination mercaptans.

[0007] The second step generally aims to minimize the amount of recombination mercaptans. The temperature is usually higher in the second step to thermodynamically promote the removal of mercaptans. In practice, a furnace is therefore placed between the two reactors to raise the temperature of the second reactor to a higher temperature than that of the first.

[0008] 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 allow the total sulfur and mercaptan content of hydrocarbon cuts, preferably gasoline cuts, to be reduced to very low levels, while minimizing the decrease in the octane rating. Typically, the catalyst used is nickel-based.

[0009] Nevertheless, there is still a need to maximize performance in the hydrotreatment of gasoline cuts to meet sulfur specifications.

[0010] Surprisingly, the Applicant has identified that a sequence of two specific catalysts of different natures, in a particular order, in the finishing hydrodesulfurization section located downstream of the selective hydrodesulfurization (SHD) section, exhibits a synergistic effect in terms of selectivity by minimizing olefin saturation, which leads to a loss in octane number. Indeed, the choice of the correct active phase and a suitable support allows for a synergy between a first catalyst in the finishing section, consisting of an active phase based on a Group VIII element, enabling the removal of recombination 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, enabling the removal of the more refractory sulfur compounds.On the other hand, the introduction of a second catalyst in the finishing section, consisting of an active phase of a group VIII element and a group VIB element, makes it possible to decrease the average treatment temperature in the HDS section and thus increase the overall cycle time of the process. Objects of the invention

[0011] The purpose of the pre-sale invention is to implement a process for producing low sulfur gasoline, making it possible to utilize the entirety of a gasoline cut containing sulfur, preferably a catalytic cracking gasoline cut, and to reduce the sulfur content in said gasoline cut to very low levels, without reducing gasoline yield while minimizing the decrease in octane rating due to olefin hydrogenation.

[0012] Thus, the present invention relates to a process for treating a gasoline containing sulfur compounds and olefins, the process comprising at least the following steps:

[0013] a) in a first reaction section, 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 sulfide form, and an oxide support, are brought into contact at a temperature between 200°C and 350°C, at a pressure between 0.2 MPa and 5 MPa, with an hourly volumetric rate between 1 h 1 and 20 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the feed flow rate to be treated expressed in m3 per hour under standard conditions between 10 Nm3 / m3 and 1000 Nm3 / m3, to obtain a first partially desulfurized effluent;

[0014] b) without separation of the H2S formed in step a), the first partially desulfurized effluent obtained at the end of step a) is directly contacted in a second reaction section with a first finishing hydrodesulfurization catalyst 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 volumetric rate between 1 h1 and 40 h1, to obtain a second partially desulfurized effluent;

[0015] c) without separation of the H2S formed in step b), the second partially desulfurized effluent obtained at the end of step b) is directly contacted in a third reaction section with a second finishing hydrodesulfurization catalyst comprising an active phase comprising, preferably made up of, 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 between 0.2 MPa and 5 MPa, with an hourly volumetric rate between 1 h1 and 40 h1, to obtain a third desulfurized effluent.

[0016] According to one or more embodiments, said second reaction section containing the first finishing hydrodesulfurization catalyst occupies a volume VI, 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 sections.

[0017] According to one or more embodiments, the catalyst of step a) and / or step c) comprises a group VIII metal content of between 0.1 and 10% by weight of group VIII metal oxide relative to the total weight of the catalyst, and a group VIB metal content of between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst.

[0018] According to one or more embodiments, the catalyst of step a) and / or of step (c) comprises alumina and an active phase comprising cobalt and molybdenum, said catalyst containing a weight content relative to the total weight of cobalt oxide catalyst, in the form of CoO, of between 0.1 and 10% and a weight content relative to the total weight of molybdenum oxide catalyst, in the form of MoO3, of between 1 and 20%, a cobalt / molybdenum molar ratio of between 0.1 and 0.8 mol / mol.

[0019] 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 catalyst of phosphorus oxide in the form of P2O5 of between 0.3 and 10% by weight.

[0020] According to one or more embodiments, the catalyst of step a) and / or step c) comprises a specific surface area between 60 and 250 m2 / g.

[0021] According to one or more embodiments, the catalysts of steps a) and c) are identical.

[0022] 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.

[0023] According to one or more embodiments, the catalyst of step b) comprises an alumina support and an active phase made of nickel, said catalyst containing a content by weight relative to the total weight of nickel oxide catalyst, in the form of NiO, of between 5 and 65% by weight.

[0024] According to one or more embodiments, the catalyst of step b) comprises a specific surface area between 60 and 250 m2 / g.

[0025] 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.

[0026] According to one or more embodiments, steps b) and c) are carried out in a single reactor.

[0027] According to one or more embodiments, the temperature of steps b) and c) is greater than the temperature of step a).

[0028] According to one or more embodiments, the gasoline is a catalytic cracking gasoline. Detailed description of the invention Definitions

[0029] In the following, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC Press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII (or VIIIB) according to the CAS classification corresponds to the metals in columns 8, 9 and 10 according to the new IUP AC classification.

[0030] The specific surface area BET is measured by nitrogen physisorption according to ASTM D3663-03, a method described in the book Rouquerol F.; Rouquerol J.; Singh K. “Adsorption by Powders & Porous Solids: Principle, methodology and applications”, Academy Press, 1999.

[0031] In the following description of the invention, the total porous volume of the oxide support or catalyst is understood to be 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 140° following the recommendations of the book "Techniques de l'ingénieur, traité analyse et caractérisation", p. 1050-5, written by Jean Charpin and Bernard Rasneur.

[0032] In order to obtain better accuracy, the value of the total pore volume in ml / g or in cmVg given in the following text corresponds to the value of the total mercury volume (total pore volume measured by intrusion in the mercury porosimeter) in ml / g or in cm 3 / g measured on the sample less the value of the mercury volume in ml / g or in cmVg measured on the same sample for a pressure corresponding to 30 psi (approximately 0.2 MPa).

[0033] The contents of group VIII metal, group VIB metal and phosphorus are measured by X-ray fluorescence.

[0034] 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. The loss on ignition is due to moisture loss. It is determined according to ASTM D7348. The charge

[0035] The process according to the invention makes it possible to treat any type of gasoline fraction containing sulfur compounds and olefins, alone or in mixtures, such as, for example, a fraction from a coking, visbreaking, steam cracking, or fluid catalytic cracking (FCC) unit. 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 fraction from a catalytic cracking unit.

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

[0037] The sulfur content of gasoline fractions produced by catalytic cracking (CC) depends on the sulfur content of the feed processed by the CC, whether or not the CC feed is pretreated, and the endpoint of the fraction. Generally, the sulfur content of an entire gasoline fraction, particularly those from CC, is greater than 100 ppm by weight and most often greater than 500 ppm by weight. For gasolines with endpoints above 200°C, the sulfur content is often greater than 1000 ppm by weight; in some cases, it can even reach values ​​of around 4000 to 5000 ppm by weight.

[0038] The feed treated by the process according to the invention can be a feed containing sulfur compounds in a content greater than 200 ppm by weight of sulfur, and often greater than 500 ppm.

[0039] Furthermore, gasoline from catalytic cracking units (FCC) contains, on average, between 0.5% and 5% weight of diolefins, between 20% and 50% weight of olefins, between 10 ppm and 0.5% weight of sulfur, of which generally less than 300 ppm of mercaptans. Step aO) Selective hydrogenation (optional)

[0040] 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 in order to hydrogenate at least partially the diolefins and carry out a weighting reaction of part of the light mercaptans (RSH) present in the thioether feed, by reaction with olefins.

[0041] To this end, the gasoline to be treated is sent to a selective hydrogenation catalytic reactor containing at least one fixed or moving bed of a catalyst for the selective hydrogenation of diolefins and the weighting of light mercaptans. The selective hydrogenation of diolefins and weighting of light mercaptans reaction 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 most preferably molybdenum.

[0042] The catalyst oxide support is preferably selected from alumina, nickel aluminate, silica, silicon carbide, or a mixture of these oxides. Alumina is preferably used, and even more preferably, high-purity alumina. In a preferred embodiment, the selective hydrogenation catalyst contains nickel with a nickel oxide content, in the form of NiO, of between 1 and 12% by weight, and molybdenum with a molybdenum oxide content, in the form of MoO3, of between 6% and 18% by weight and a nickel / molybdenum molar ratio of between 0.3 and 2.5, the metals being deposited on an alumina support. The degree of sulfidation of the metals constituting the catalyst is preferably greater than 60%.

[0043] During the optional selective hydrogenation step, the gasoline is brought into contact with the catalyst at a temperature between 50 and 250°C, and preferably between 80 and 220°C, and even more preferably between 90 and 200°C, with a volumetric flow rate per hour (WH) between 0.5 h⁻¹ and 20 h⁻¹, the unit of the volumetric flow rate per hour being the volumetric charge flow rate 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 m3 per hour and the volumetric flow rate of feed to be treated expressed in m3 per hour under standard conditions (15°C, 0.1 MPa) of between 2 and 100 Nm3 / m3, preferably between 3 and 30 Nm3 / m3.

[0044] 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, it can be obtained to less than 1 mg AM / g.

[0045] 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 two-cut fractionation, the heavy cut is processed according to the method of the invention. In the case of three-cut fractionation, the intermediate and heavy cuts can be processed separately according to the method of the invention.

[0046] It should be noted that it is possible to carry out the hydrogenation steps of the diolefins and the fractionation into two or three cuts simultaneously using a catalytic distillation column which includes a distillation column equipped with minus a catalytic bed. Step a) Selective hydrodesulfurization (HDS)

[0047] The hydrodesulfurization step a) is implemented to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H2S.

[0048] 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.

[0049] The operating pressure of this step is generally between 0.2 MPa and 5 MPa, preferably between 1 MPa and 3 MPa.

[0050] The quantity of catalyst used in each reactor of the first reaction section is generally such that the ratio between the volumetric flow rate at 15°C of gasoline to be treated, expressed in m3 per hour, per m3 of catalytic bed (also called hourly volumetric rate - WH) is between 1 and 20 h1 and preferably between 2 and 10 h1.

[0051] The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m3 per hour (Nm3 / h) and the volumetric flow rate of the feed to be treated expressed in m3 per hour under standard conditions (15°C, 0.1 MPa) is between 10 and 1000 Nm3 / m3, preferably between 50 and 600 Nm3 / m3. Normal m3 refers to the volume of 1 m3 of gas at 0°C and 0.1 MPa.

[0052] The hydrogen required for this step can be fresh hydrogen or recycled hydrogen, preferably free of H2S, or a mixture of fresh and recycled hydrogen. Preferably, a mixture of fresh and recycled hydrogen will be used.

[0053] 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 weight of sulfur and preferably less than 100 ppm weight of sulfur.

[0054] 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.

[0055] According to the invention, the hydrodesulfurization catalyst of step a) comprises an active phase comprising, preferably made up 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.

[0056] The VIB group metal present in the active phase of the catalyst is preferably chosen from molybdenum and tungsten.

[0057] The group VIII metal present in the active phase of the catalyst is preferably chosen from cobalt, nickel and a mixture of these two elements.

[0058] 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 most preferably the active phase consists of cobalt and molybdenum.

[0059] The Group VIII metal content 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 most 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.

[0060] The metal content of group VIB is preferably between 1 and 20% by weight of the oxide of the group VIB metal relative to the total weight of the catalyst, more preferably between 2 and 18% by weight, and most preferably between 3 and 16% by weight of the oxide of the group VIB metal relative to the total weight of the catalyst. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 or WO3.

[0061] 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.

[0062] 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, most preferably between 0.5 and 3% by weight.

[0063] Moreover, when phosphorus is present, the phosphorus / (metal of group VIB) molar ratio is generally between 0.1 and 0.7 mol / mol, preferably between 0.2 and 0.6 mol / mol.

[0064] Preferably, the catalyst of step a) comprises a specific surface area of ​​between 60 and 250 m2 / g, preferably between 60 and 200 m2 / g, and even more preferably between 65 and 180 m2 / g, and even more preferably between 70 and 130 m2 / g.

[0065] The total porous volume of the catalyst in step a) is generally between 0.3 cmVg and 1.3 cmVg, preferably between 0.4 cmVg and 1.1 cmVg.

[0066] The oxide support for the hydrodesulfurization catalyst is typically a porous solid selected from the group consisting of: alumina, 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. Most preferably, the oxide support is essentially composed of alumina, that is to say, it comprises at least 51% weight, preferably at least 60% by weight, most 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.

[0067] In a preferred embodiment, the catalyst of step a) comprises an alumina support and an active phase comprising, preferably composed of, cobalt and molybdenum and optionally phosphorus, said catalyst containing a weight content relative to the total weight of cobalt oxide catalyst, in the form of CoO, of between 0.1 and 10 wt%, preferably between 0.6 and 8 wt%, more preferably between 0.6 and 7 wt% and even more preferably between 1 and 6 wt%, and a weight content relative to the total weight of molybdenum oxide catalyst, in the form of MoO3, of between 1 and 20 wt%, preferably between 2 and 18 wt%, and most preferably between 3 and 16 wt%, with a cobalt / molybdenum molar ratio of between 0.1 and 0.8 mol / mol, preferably between 0.2 and 0.6 mol / mol.

[0068] Preferably, the hydrodesulfurization catalyst support comprises a specific surface area of ​​between 60 and 250 m2 / g, preferably between 60 and 200 m2 / g, and even more preferably between 65 and 180 m2 / g, and even more preferably between 70 and 130 m2 / g.

[0069] The total porous volume of the hydrodesulfurization catalyst support is generally between 0.3 cmVg and 1.3 cmVg, preferably between 0.4 cmVg and 1.1 cmVg.

[0070] The hydrodesulfurization catalyst support may be in the form of beads, extrudates of any geometry, wafers, 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, trilobed, or quadrilobed extrudates with a circumscribed diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0071] 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.

[0072] Step b) First finishing hydrodesulfurization step (FNS1)

[0073] During the hydrodesulfurization step a), a large part of the sulfur compounds are transformed into H2S. The remaining sulfur compounds are essentially refractory sulfur compounds and recombination mercaptans resulting from the addition of the H2S formed in step a) to the olefins present in the feed.

[0074] Step b) of the process according to the invention consists of transforming at least a portion of the recombination mercaptans contained in the first effluent from step a) into olefins and H2S, as well as at least a portion of the sulfur compounds contained in the first effluent from step a), such as thiophenic compounds, into compounds saturated, for example, with thiophanes (or thiacyclopentanes) or mercaptans, and then hy- to at least partially degrade these sulfur compounds to form H2S.

[0075] Preferably, step b) is carried out at a higher temperature than step a). Indeed, using a higher temperature in this step compared to the temperature of step a) will discourage the formation of mercaptans by shifting the thermodynamic equilibrium. Step b) also allows the hydrodesulfurization of residual sulfur compounds to continue.

[0076] 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.

[0077] The operating pressure of this step is generally between 0.2 MPa and 5 MPa and preferably between 1.5 MPa and 3 MPa.

[0078] The quantity of catalyst used in each reactor of the second reaction section is generally such that the ratio between the volumetric flow rate of gasoline to be treated, expressed in m3 per hour under standard conditions (15°C, 0.1 MPa), per m3 of catalytic bed (also called hourly volumetric rate - WH) is between 1 and 40 h1 and preferably between 2 and 20 h1.

[0079] The first hydrodesulfurization 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 sulfided.

[0080] The Group VIII metal is preferably nickel. When the Group VIII metal is nickel, the phase diagram of nickel sulfide exhibits a large number of sulfur-rich and nickel-rich phases at low temperatures. Various phases and stoichiometries of nickel sulfide are therefore possible, ranging from nickel-rich compounds such as Ni3S2, Ni6S5, Ni7S6, Ni9S8, and NiS to sulfur-rich compounds such as Ni3S4 and NiS2. It should be noted that NiS is also known to exist in two main phases: hexagonal α-NiS, stable at high temperatures, and rhombohedral α-NiS, stable at low temperatures. The existence of these numerous phases makes the synthesis of nickel sulfide as a single phase complex; the products are therefore often mixtures of two or more phases.

[0081] The Group VIII metal content is preferably between 5 and 65% by weight of Group VIII metal oxide relative to the total weight of the catalyst, more preferably between 8 and 55% by weight, and most 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.

[0082] Preferably, the catalyst of step b) is characterized by a specific surface area of ​​between 60 and 250 m2 / g, preferably between 70 and 200 m2 / g.

[0083] The total porous volume of the catalyst in step b) is generally between 0.3 cm3 / g and 1.3 cm3 / g, preferably between 0.4 cm3 / g and 1.1 cm3 / g.

[0084] The oxide support for the first finishing hydrodesulfurization catalyst is typically a porous solid selected from the group consisting of: alumina, 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. Most preferably, the oxide support is essentially composed of alumina, that is to say, it comprises at least 51% by weight, preferably at least 60% by weight, most 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.

[0085] Preferably, the support for the first finishing hydrodesulfurization catalyst comprises a specific surface area of ​​between 60 and 250 m2 / g, preferably between 70 and 200 m2 / g.

[0086] The total porous volume of the support for the first finishing hydrodesulfurization catalyst is generally between 0.3 cmVg and 1.3 cmVg, preferably between 0.4 cmVg and 1.1 cmVg.

[0087] The support for the first finishing hydrodesulfurization catalyst may be in the form of beads, extrudates of any geometry, wafers, 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, trilobed, or quadrilobed extrudates with a circumscribed diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0088] 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.

[0089] Step c) Second finishing hydrodesulfurization step (FNS2)

[0090] The sulfur compounds remaining after step b) are essentially refractory sulfur compounds. Step c) of the process according to the invention essentially consists of transforming at least a part of the refractory sulfur compounds contained in the effluent from step b) such as thiophenic compounds, into saturated compounds for example in thiophanes (or thiacyclopentanes) or in mercaptans, and then at least partially hydrogenolyzing these sulfur compounds to form H2S.

[0091] 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.

[0092] The operating pressure of this step is generally between 0.2 MPa and 5 MPa and preferably between 1.5 and 3 MPa.

[0093] The quantity of catalyst used in each reactor is generally such that the ratio between the volumetric flow rate of gasoline to be treated, expressed in m3 per hour, and Under standard conditions (15°C, 0.1 MPa), the volumetric rate per m³ of catalytic bed (also called hourly volumetric rate) is between 1 and 40 h₁ and preferably between 2 and 20 h₂

[0094] In the process according to the invention, the hydrogenation rate of the olefins in steps b) and c) is preferably less than 30% during these steps.

[0095] 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 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.

[0096] The second finishing hydrodesulfurization catalyst of step c) comprises an active phase comprising, preferably made up of, at least one metal from group VIB and at least one metal from group VIII at least partly in sulfide form, optionally phosphorus, and an oxide support, as described below.

[0097] The metal from group VIB is preferably chosen from molybdenum and tungsten. The metal from group VIII is preferably chosen from cobalt, nickel, and mixtures of these two elements. The active phase of the catalyst is preferably chosen from the group formed by the combination of nickel-molybdenum, cobalt-molybdenum, and nickel-cobalt-molybdenum, and most preferably the active phase consists of cobalt and molybdenum.

[0098] The Group VIII metal content is preferably between 0.1 and 10% by weight of the 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 most preferably between 1 and 6% by weight. When the metal is cobalt or nickel, the metal content is expressed as CoO or NiO.

[0099] The metal content of group VIB is preferably between 1 and 20% by weight of the oxide of the group VIB metal relative to the total weight of the catalyst, more preferably between 2 and 18% by weight, and most preferably between 3 and 16% by weight. When the metal is molybdenum or tungsten, the metal content is expressed in MoO3 or WO3.

[0100] 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.

[0101] Optionally, the catalyst may also have a phosphorus content generally between 0.3 and 10 wt% of P2O5 relative to the total weight of catalyst, preferably between 0.3 and 5 wt%, most preferably between 0.5 and 3 wt%. Furthermore, when phosphorus is present, the molar ratio phosphorus / (metal of group VIB) is generally between 0.1 and 0.7 mol / mol, preferably between 0.2 and 0.6 mol / mol.

[0102] Preferably, the catalyst of step c) comprises a specific surface area of ​​between 60 and 250 m2 / g, preferably between 60 and 200 m2 / g, and even more preferably between 65 and 180 m2 / g, and even more preferably between 70 and 130 m2 / g.

[0103] The total porous volume of the catalyst in step c) is generally between 0.3 cmVg and 1.3 cmVg, preferably between 0.4 cmVg and 1.1 cmVg.

[0104] The oxide support for the hydrodesulfurization catalyst is typically a porous solid selected from the group consisting of: alumina, 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. Most preferably, the oxide support is essentially composed of alumina, that is to say, it comprises at least 51% by weight, preferably at least 60% by weight, most 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.

[0105] In a preferred embodiment, the catalyst of step c) comprises an alumina support and an active phase comprising, preferably, cobalt and molybdenum, said catalyst containing a weight content relative to the total weight of cobalt oxide catalyst, in the form of CoO, of between 0.1 and 10 wt%, preferably between 0.6 and 8 wt%, more preferably between 0.6 and 7 wt%, and even more preferably between 1 and 6 wt%, and a weight content relative to the total weight of molybdenum oxide catalyst, in the form of MoO3, of between 1 and 20 wt%, preferably between 2 and 18 wt%, and most preferably between 3 and 16 wt%, with a cobalt / molybdenum molar ratio of between 0.1 and 0.8 mol / mol, preferably between 0.2 and 0.6 mol / mol.

[0106] Preferably, the support for the second finishing hydrodesulfurization catalyst comprises a specific surface area of ​​between 60 and 250 m2 / g, preferably between 60 and 200 m2 / g, and even more preferably between 65 and 180 m2 / g, and even more preferably between 70 and 130 m2 / g.

[0107] The total porous volume of the support for the second finishing hydrodesulfurization catalyst is generally between 0.3 cmVg and 1.3 cmVg, preferably between 0.4 cmVg and 1.1 cmVg.

[0108] The support for the second finishing hydrodesulfurization catalyst may be in the form of beads, extrudates of any geometry, wafers, 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, trilobed or quadrilobed extrades with a circumscribed diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0109] In a preferred embodiment, the catalyst implemented in step c) is the same as that implemented in step a).

[0110] Preferably, step c) is carried out in the same reactor as step b). Implementation of steps b) and c) of the process

[0111] Steps b) and c) of the process according to the invention can be carried out in one, two or more reactors.

[0112] 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 from step a passes, and then step c) can be carried out in the second finishing hydrodesulfurization reactor containing the third reaction section, placed downstream of said first reactor.

[0113] 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.

[0114] According to one or more embodiments, said second reaction section containing the first finishing hydrodesulfurization catalyst occupies a volume VI, 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 hydrodesulfurization reaction sections, preferably between 90%vol / 10%vol and 60%vol / 40%vol respectively of said second and third finishing hydrodesulfurization reaction sections. Step d): Separation of H2S (optional)

[0115] The separation step d) is implemented in order to separate the excess hydrogen as well as the H2S formed during steps a), b) and c). Any method known to a person skilled in the art may be considered.

[0116] 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 to condense the hydrocarbons. The gas and liquid phases are then separated in a separation vessel. The liquid fraction, which contains the desulfurized gasoline and a fraction of the dissolved H2S, is sent to a stabilization or starter column. This column separates an essentially consisting of residual H2S and hydrocarbon compounds having a boiling point lower than or equal to that of butane and a bottom cut free of H2S, called stabilized gasoline, containing compounds having a boiling point higher than that of n-butane.

[0117] According to a second embodiment, after the condensation step, the liquid fraction containing the desulfurized gasoline and a fraction of the dissolved H2S is sent to a stripping section, while the gaseous fraction, consisting mainly of hydrogen and H2S, is sent to a purification section. Stripping can be carried out by heating the hydrocarbon fraction alone, or with the injection of hydrogen or steam, in a distillation column to extract, at the top, the light compounds that have been carried along by dissolution in the liquid fraction, as well as the residual dissolved H2S. The temperature of the stripped gasoline recovered at the bottom of the column is generally between 120°C and 250°C.

[0118] Preferably, the separation step d) is carried out in a stabilization or starter column. Indeed, a stabilization column allows H2S to be separated more efficiently than a stripping section.

[0119] Step d) is preferably implemented so that the sulfur in the form of H2S 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. Catalyst preparation

[0120] 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 impregnating the selected porous support with elements from groups VIII and possibly VIB and phosphorus. Impregnation can, for example, be carried out using the method known to those skilled in the art, referred to as dry impregnation, in which only 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 precisely as possible. Preferably, the aqueous impregnation solution, when it contains cobalt, molybdenum, and phosphorus, is prepared under pH conditions that favor the formation of heteropolyanions in solution. For example, the pH of such an aqueous solution is between 1 and 5.Preferably, the catalyst preparation is carried out without the addition of any organic agent mixed with the precursors of group VIII, group VI and phosphorus elements.

[0121] 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, phosphomo- Molybdenum trioxide (H3PMoi2O4O), and their salts, and possibly silicomolybdic acid (H4SiMoi2O4O) and its salts, can be used as sources. Sources of molybdenum can also include any heteropolycompound of the Keggin, lacunar Keggin, substituted Keggin, Dawson, Anderson, and Strandberg types, for example. Molybdenum trioxide and heteropolycompounds of the Keggin, lacunar Keggin, substituted Keggin, and Strandberg types are preferred.

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

[0123] The nickel precursors that can be used are advantageously chosen from among oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example. Nickel hydroxide and nickel hydroxycarbonate are preferred.

[0124] The tungsten precursors that can be used are also well known to those skilled in the art. For example, tungsten sources include oxides and hydroxides, tungstic acids and their salts, in particular ammonium salts such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and their salts, and possibly silicotungstic acid (H4SiWi2O4O) and its salts. Tungsten sources can also be any heteropolycompound of the Keggin, lacunar Keggin, substituted Keggin, or Dawson type, for example. Ammonium oxides and salts such as ammonium metatungstate or heteropolyanions of the Keggin, lacunar Keggin, or substituted Keggin type are preferred.

[0125] Phosphorus can advantageously be introduced alone or in a mixture with at least one of the elements of Group VIB and Group VIII. Preferably, phosphorus is introduced in a mixture with the precursors of the Group VIB and Group VIII elements by dry impregnation of said porous support with a solution containing the element precursors 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 simultaneously with the Group VIB element(s) in the form of, for example, Keggin, lacunar Keggin, substituted Keggin, or Strandberg-type heteropolyanions.

[0126] 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.

[0127] Following the maturation step, the resulting catalyst precursor 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 a oxidative treatment but simple drying of the catalyst can also be carried out.

[0128] In the case of drying, the catalyst precursor is dried at a temperature between 50°C and below 200°C, preferably between 70°C and 180°C, for a period typically between 0.5 hours and 12 hours, and even more preferably for a period between 0.5 hours and 5 hours.

[0129] In the case of an oxidizing treatment, also called calcination, this is generally carried out under air or dilute oxygen, and the treatment temperature is generally between 200°C and 550°C, preferably between 300°C and 500°C, and advantageously for a period typically between 0.5 hours and 24 hours, preferably for a period of 0.5 hours to 12 hours, and even more preferably for a period of 0.5 hours to 10 hours.

[0130] Before its use as a hydrotreating catalyst, it is advantageous to subject the optionally dried or calcined catalyst to a sulfidation activation step. This activation phase is carried out by methods well known to those skilled in the art, and advantageously under a sulfur-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 the sulfidation of catalysts

[0131] Before contacting the feedstock in a gasoline hydrodesulfurization process, the catalysts used in steps (optionally a0), a), b), and c) of the process according to the invention generally undergo a sulfidation step. Sulfuration is preferably carried out in a sulfur-reducing medium, i.e., in the presence of H2S and hydrogen, in order to transform metal oxides into sulfides such as, for example, MoS2, Co9S8, or Ni3S2. Sulfuration is performed by injecting a stream containing H2S and hydrogen, or a sulfur compound capable of decomposing into H2S in the presence of the catalyst and hydrogen, onto the catalyst. Polysulfides such as dimethyl disulfide (DMDS) are H2S precursors commonly used to sulfidize the catalysts in steps (optionally a0), a), b), and c). The sulfur can also originate from the feedstock.The temperature is adjusted so that H2S reacts with metal oxides to form metal sulfides. This sulfidation can be carried out in situ or ex situ (inside or outside the reactor) of the reactor of the process according to the invention at temperatures between 200°C and 600°C, and more preferably between 300°C and 500°C.

[0132] The degree of sulfidation of the metals constituting the catalysts of the steps, optionally aO), a), b), or c), is at least 60%, preferably at least 70%. The sulfur content in the sulfided catalyst of the steps, optionally aO), a), b), or c), is measured by elemental analysis according to ASTM D5373. A metal is considered sulfided when the overall degree of sulfidation defined by the ratio The molar ratio between sulfur (S) present on the catalyst and the metal in question is at least equal to 60% of the theoretical molar ratio corresponding to the total sulfidation of the metal(s) considered. The overall sulfidation rate is defined by the following equation:

[0133] (S / metal)cataiyseur > 0.6 x (S / metal)theoretical

[0134] in which:

[0135] (S / metal)catalyst is the molar ratio between sulfur (S) and metal present on the catalyst

[0136] (S / metal)theoretical is the molar ratio between sulfur and metal corresponding to the total sulfidation of the metal into sulfide.

[0137] This theoretical molar ratio varies depending on the metal considered:

[0138] - (S / Co)theoretical= 1

[0139] - (S / Ni)theoretical = 1

[0140] - (S / Mo)theoretical=2 / 1

[0141] - (S / W)theoretical=2 / l

[0142] When the catalyst used in step (optionally a0), 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 into sulfide, the calculation being carried out in proportion to the relative mole fractions of each metal.

[0143] The following examples illustrate the invention without limiting its scope. Examples

[0144] The analytical methods used to characterize the loads and effluents are as follows:

[0145] - sulfur content according to ASTM D2622 method for contents greater than 10 ppm S and ISO 20846 for levels below 10 ppm S;

[0146] - mercaptan content according to ASTM D3227 method;

[0147] - olefin content based on gas chromatography analysis according to the ASTM D6733 method. Example 1: Preparation of catalyst A

[0148] A support A' composed of alumina in the form of beads with a grain size between 2 and 4 mm, and having a specific surface area of ​​139 m2 / g and a pore volume of 0.97 mL / g, is provided.

[0149] Cobalt and molybdenum are then added. The impregnation solution is prepared by dissolving ammonium heptamolybdate tetrahydrate (5.64 g, >99.5%, Sigma-Aldrich®) and cobalt nitrate hexahydrate (5.36 g, >99.5%, Alfa Aesar®) at room temperature in 28 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, and finally calcined under an air flow of IL / h / g at 450°C for 4 hours. The catalyst thus obtained is denoted A.

[0150] The final composition in elements of catalyst A, expressed in the form of oxides, and referred to the weight of the dry catalyst is then as follows: MoO3 = 10.0 + / - 0.2 wt% and CoO = 3.0 + / - 0.1 wt%.

[0151] The molar ratio Co / Mo is 0.60 mol / mol.

[0152] The specific surface area of ​​catalyst A is 124 m2 / g. Example 2: Preparation of catalyst B

[0153] A support B' identical to the support A' is provided.

[0154] 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 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 of IL / h / g at 450°C for 4 hours. The catalyst thus obtained is denoted B.

[0155] The final composition in elements of catalyst B, expressed in the form of oxides and referred to the weight of the dry catalyst is then as follows: NiO = 17.9 + / - 0.3 wt.

[0156] The specific surface area of ​​catalyst B is 114 m2 / g.

[0157] Example 3: Implementation of catalysts in a gasoline desulfurization process

[0158] Example 3 aims to demonstrate the advantages of the gasoline desulfurization process using a sequence of steps and specific catalysts for each step. Gasoline from a catalytic cracking unit, composed of 25% olefins by weight and 600 ppmS of total sulfur, is subjected to a multi-step treatment:

[0159] - a selective hydrodesulfurization (SDS) step in an adiabatic reactor putting in The catalyst A is used. The operating conditions of the one-stage hydrodesulfurization step of the gasoline feed are as follows: WH = 3 h*, P = 2.0 MPa. A stream of pure hydrogen is added to the reactor inlet feed such that H2 / HC = 250 Nm3 / m3. The effluent is sent directly to the second-stage reactor;

[0160] - a first finishing hydrodesulfurization step (FNS1) in an adiabatic reactor using catalysts A or B. Only the effluent from the first stage is treated in this second stage. The pressure of the first finishing hydrodesulfurization stage is set at 2.0 MPa. The reactor inlet temperature is always set at 35°C higher than the temperature of the first effluent. exit from the selective hydrodesulfurization stage;

[0161] - possibly, a second finishing hydrodesulfurization step (FNS2) in Adiabatic reactor using catalysts A or B. Only the effluent from the previous stage is treated in this second, final hydrodesulfurization stage. The pressure of the second final hydrodesulfurization stage is set at 2.0 MPa. The reactor inlet temperature is equal to the effluent outlet temperature of the previous stage.

[0162] The inlet temperature to the selective hydrodesulfurization stage reactor is set to obtain an effluent containing 10 ppm wt S of total sulfur (i.e., more than 98% conversion to total sulfur). Prior to use, the catalysts contained in the selective hydrodesulfurization and finishing reactors are sulfided by treatment for 4 hours under a pressure of 3.4 MPa at 350°C, in contact with a feed consisting of 2 wt% sulfur in the form of dimethyl disulfide (DMDS) in n-heptane.

[0163] The results illustrate that the gasoline hydrodesulfurization process according to the invention provides the best performance compared to known prior art implementations, 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 HDS section, thereby increasing the lifespan of the catalysts. The performance of the gasoline desulfurization process is presented in Table 1.

[0164] [Tables 1] Implementation Non-compliant implementation Non-compliant implementation Non-compliant implementation Implementation compliant with the invention HDS Catalyst AAAA WH (h') 3 3 3 3 Reactor inlet temperature (°C) 253 220 240 237 Reactor outlet temperature (°C) 289 224 260 255 Average temperature (°C) 271 222 250 246 FNS1 Catalyst BAAB WH (h') 3 3 12 4 Reactor inlet temperature (°C) 324 259 295 290 Reactor outlet temperature (°C) 326 294 317 291 Average temperature (°C) 325 277 306 291 FNS2 Catalyst - - BA WH (h') - - 4 12 Reactor inlet temperature (°C) - - 317 291 Temperature Reactor outlet (°C) - - 319 310 Average temperature (°C) 318 301 Overall performance FNS1 / FNS2 ratio (%vol) - - 25 / 75 75 / 25 Inlet olefins (%wt) 25.0 25.0 25.0 25.0 Inlet olefins (%wt) 16.5 15.6 15.3 16.7 Overall olefin hydrogenation (%) 34.0 37.6 38.8 33.2 Inlet sulfur content (ppm wt S) 600 600 600 600 Sulfur content at outlet (ppm wt S) 10 10 10 10

[0165] The "average temperature" of the HDS, FNS1, or FNS2 stage corresponds to the Weight Average Bed Temperature (WABT), a term well known to those skilled in the art. The average temperature is advantageously determined based on the catalytic systems, equipment, and their configuration used. The average temperature (or WABT) is calculated as follows:

[0166] [Math.l] WAH i

[0167] with Tinlet: the temperature of the flow at the inlet of the reaction section and Toutlet: the temperature of the effluent at the outlet of the reaction section. Unless otherwise specified, the "average temperature" of a reaction section is given at the start of the cycle.

Claims

Demands

1. Process for treating gasoline containing sulfur compounds and olefins, the process comprising at least the following steps: a) in a first reaction section, the gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a metal of group VIB and a metal of group VIII at least partly in sulfide form, and an oxide support are brought into contact, at a temperature between 200°C and 350°C, at a pressure between 0.2 MPa and 5 MPa, with an hourly volumetric rate between 1 h 1 and 20 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the volumetric flow rate of feed to be treated expressed in m3 per hour under standard conditions between 10 Nm3 / m3 and 1000 Nm3 / m3, 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 contacted in a second reaction section with a first finishing hydrodesulfurization catalyst 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 volumetric rate between 1 h1 and 40 h1, 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 contacted in a third reaction section with a second finishing hydrodesulfurization 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 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 volumetric rate between 1 h⁻¹ and 40 h⁻¹*, to obtain a third desulfurized effluent.

2. The process according to claim 1, wherein said second reaction section containing the first finishing hydrodesulfurization catalyst occupies a volume VI, and said third finishing hydrodesulfurization reaction section containing the second catalyst the finishing hydrodesulfurization 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. A process according to any one of claims 1 or 2, wherein the catalyst of step a) and / or step c) comprises a Group VIII metal content of between 0.1 and 10% by weight of Group VIII metal oxide relative to the total weight of the catalyst, and a Group VIB metal content of between 1 and 20% by weight of Group VIB metal oxide relative to the total weight of the catalyst.

4. A process according to any one of the preceding claims, wherein the catalyst of step a) and / or step c) comprises alumina and an active phase comprising cobalt and molybdenum, said catalyst containing a weight relative to the total weight of cobalt oxide catalyst, in the form of CoO, of between 0.1 and 10% and a weight relative to the total weight of molybdenum oxide catalyst, in the form of MoO3, of between 1 and 20%, a cobalt / molybdenum molar ratio of between 0.1 and 0.8 mol / mol.

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

6. A process according to any one of the preceding claims, wherein the catalyst of step a) and / or step c) comprises a specific surface area between 60 and 250 m2 / g.

7. A method according to any one of the preceding claims, wherein the catalysts of steps a) and c) are identical.

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

10. A method according to any one of the preceding claims, in which the catalyst of step b) comprises a specific surface area between 60 and 250 m2 / g.

11. A method according to any one of the preceding claims, wherein prior to step a), the gasoline is contacted with hydrogen and a selective hydrogenation catalyst to selectively hydrogenate the diolefins contained in said gasoline into olefins.

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

13. A method according to any one of the preceding claims, wherein the temperature of steps b) and c) is greater than the temperature of step a).

14. A method according to any one of the preceding claims, wherein the gasoline is a catalytic cracking gasoline.