Process for finishing gasoline hydrodesulfurization using catalyst chain

Abstract

A process for treating gasoline containing sulfur compounds and olefins, comprising the steps of: a) contacting gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a Group VIB metal and a Group VIII metal with a support in a first reaction section, thereby obtaining a first partial desulfurization effluent; b) in a second reaction section, contacting the first partial desulfurization effluent with a first refined hydrodesulfurization catalyst comprising an active phase consisting of a Group VIII metal and a support, thereby obtaining a second partial desulfurization effluent; c) in a third reaction section, contacting the second partial desulfurization effluent with a second refined hydrodesulfurization catalyst comprising an active phase and a support, the active phase comprising a Group VIB metal and a Group VIII metal, thereby obtaining a third partial desulfurization effluent.

Description

Field of the Invention

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

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

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

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

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

[0006] The gasoline can thus be treated by a sequence of two reactors as described in document EP 1 077 247; the purpose of the first stage (also called the selective HDS stage) is generally to carry out a deep desulfurization of the gasoline with a minimum of olefin saturation (and without loss of aromatics), thus achieving the maximum octane retention rate. The catalyst used is generally a CoMo type catalyst. During this stage, new sulfur compounds are formed by recombination of the H2S from the desulfurization with the olefins: recombined mercaptans.

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

[0008] The catalyst used in the refining process must be particularly selective so as not to cause olefin saturation (and not cause loss of aromatics), which would lead to octane loss. Thus, it must make it possible to reduce the content of total sulfur and mercaptans in the hydrocarbon fraction, preferably the gasoline fraction, to extremely low levels while minimizing the reduction of the octane number. Generally, the catalyst used is nickel-based.

[0009] However, there is still a need to maximize the performance of the hydrotreated gasoline fraction to meet the sulfur specifications.

[0010] Surprisingly, the applicant has confirmed that in the finishing hydrodesulfurization section located downstream of the selective hydrodesulfurization (HDS) section, the sequential use of two specific catalysts with different properties in a specific order has a synergistic effect in terms of selectivity, while minimizing the saturation of olefins that cause octane number loss. Specifically, the selection of the correct active phase and a suitable support enables the observation of a synergistic effect between the first catalyst in the finishing section and the second catalyst in the finishing section. The first catalyst consists of an active phase based on Group VIII elements and is capable of removing heavy mercaptans while retaining olefins. The second catalyst consists of an active phase based on Group VIII elements and Group VIB elements and is capable of removing more refractory sulfur compounds. In addition, when the second catalyst consisting of an active phase based on Group VIII elements and Group VIB elements is introduced into the finishing section, the average treatment temperature of the HDS section can be reduced, thereby increasing the overall cycle time of the process. Summary of the Invention

[0011] The object of the present invention is to implement a method for preparing gasoline with a low sulfur content, which is capable of upgrading the entire sulfur-containing gasoline fraction, preferably the catalytic cracking gasoline fraction, and reducing the sulfur content in the gasoline fraction to a very low level without reducing the gasoline yield, while minimizing the octane number reduction caused by olefin hydrogenation.

[0012] Therefore, the subject of the present invention is a method for treating gasoline containing sulfur compounds and olefins, the method comprising at least the following steps:

[0013] a) In a first reaction section, contacting gasoline, hydrogen, and a hydrodesulfurization catalyst at a temperature of 200 °C to 350 °C and a pressure of 0.2 MPa to 5 MPa, wherein the space velocity is 1 h -1 to 20 h -1 , and the hydrogen flow rate expressed as standard m 3 / h to the flow rate of the feedstock to be treated expressed as m 3 / h under standard conditions is in the ratio of 10 Nm 3 / m 3 to 1000 Nm 3 / m 3 , and the hydrodesulfurization catalyst comprises an active phase and an oxide support, and the active phase comprises a Group VIB metal and a Group VIII metal at least partially in the sulfide form, thereby obtaining a first partially desulfurized effluent;

[0014] b) Without separating the H2S formed in step a), in a second reaction section, directly contacting the first partially desulfurized effluent obtained at the end of step a) with a first finishing hydrodesulfurization catalyst at a temperature of 250 °C to 400 °C and a pressure of 0.2 MPa to 5 MPa, wherein the space velocity is 1 h-1 to 40 h -1 , the first hydrodesulfurization catalyst comprises an active phase and an oxide support, the active phase consists of Group VIII metals at least partially in the form of sulfides, thereby obtaining a second part of the desulfurized effluent;

[0015] c) Without separating the H2S formed in step b), in the third reaction section, the second part of the desulfurized effluent obtained at the end of step b) is directly contacted with a second hydrodesulfurization catalyst at a temperature of 250 °C to 400 °C and a pressure of 0.2 MPa to 5 MPa, wherein the space velocity is 1 h -1 to 40 h -1 , the second hydrodesulfurization catalyst comprises an active phase and an oxide support, the active phase contains at least one Group VIB metal and at least one Group VIII metal at least partially in the form of sulfides, preferably consists of at least one Group VIB metal and at least one Group VIII metal at least partially in the form of sulfides, thereby obtaining a third part of the desulfurized effluent.

[0016] According to one or more embodiments, the second reaction section containing the first hydrodesulfurization catalyst occupies a volume V1, and the third hydrodesulfurization reaction section containing the second hydrodesulfurization catalyst occupies a volume V2, and the volume distribution of V1 / V2 of the second and third hydrodesulfurization reaction sections is 90 vol% / 10 vol% to 10 vol% / 90 vol%, preferably 90 vol% / 10 vol% to 60 vol% / 40 vol%.

[0017] According to one or more embodiments, the catalyst in step a) and / or step c) contains a Group VIII metal content of 0.1 wt% to 10 wt% based on the total weight of the catalyst in terms of the oxide of the Group VIII metal, and a Group VIB metal content of 1 wt% to 20 wt% based on the total weight of the catalyst in terms of the oxide of the Group VIB metal.

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

[0019] According to one or more embodiments, the catalyst of step a) and / or step c) further comprises phosphorus, and the catalyst contains 0.3 wt% to 10 wt% of phosphorus oxide in the form of P2O5 relative to the total weight of the catalyst.

[0020] According to one or more embodiments, the catalyst of step a) and / or step c) has a specific surface area of 60 to 250 m 2 / g.

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

[0022] According to one or more embodiments, the catalyst of step b) contains a Group VIII metal content of 5 wt% to 65 wt% relative to the total weight of the catalyst in terms of the oxide of the Group VIII metal.

[0023] According to one or more embodiments, the catalyst of step b) comprises an alumina support and an active phase composed of nickel, and the catalyst contains 5 wt% to 65 wt% of nickel oxide in the form of NiO relative to the total weight of the catalyst.

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

[0025] According to one or more embodiments, before step a), gasoline is contacted with hydrogen and a selective hydrogenation catalyst to selectively hydrogenate the diolefins contained in the gasoline, thereby obtaining olefins.

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

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

[0028] According to one or more embodiments, the gasoline is catalytically cracked gasoline. Detailed Description

[0029] Definition

[0030] Hereinafter, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, edited by D.R. Lide, 81st edition, 2000 - 2001). For example, Group VIII (or Group VIIIB) according to the CAS classification corresponds to the metals in columns 8, 9, and 10 according to the new IUPAC classification.

[0031] The BET specific surface area is measured by nitrogen physical adsorption according to the standard ASTM D3663-03, the method of which is described in the work “Adsorption by Powders & Porous Solids: Principles, Methodology and Applications” by Rouquerol F., Rouquerol J. and Singh K., Academic Press, 1999.

[0032] In the following description of the present invention, the “total pore volume” of the oxide support or catalyst is understood to mean the volume measured by mercury porosimetry according to the standard ASTM D4284-83 at a maximum pressure of 4000 bar (400 MPa), using a surface tension of 484 dynes / cm and a contact angle of 140°. The wetting angle is taken to be 140° according to the recommendation on page 1050-5 of the publication “Techniques de l'ingénieur, traité analyse et caractérisation” [Techniques of the Engineer, Analysis and Characterization Treatise] written by Jean Charpin and Bernard Rasneur.

[0033] For better accuracy, the values of the total pore volume given hereinafter in ml / g or in cm 3 / g correspond to the total mercury volume value measured in ml / g or in cm 3 / g on the said sample minus the mercury volume value measured in ml / g or in cm 3 / g on the same sample at a pressure corresponding to 30 psi (about 0.2 MPa).

[0034] The contents of Group VIII metals, Group VIB metals and phosphorus are measured by X-ray fluorescence.

[0035] The contents of Group VIB metals, Group VIII metals and phosphorus in the said catalyst are expressed as oxides after correcting for the loss on ignition of the catalyst sample in a muffle furnace at 550 °C for 2 h. The loss on ignition is due to the loss of moisture. It is determined according to ASTM D7348.

[0036] Raw materials

[0037] The method according to the invention enables any type of gasoline fraction containing sulfur compounds and olefins, either alone or as a mixture, such as fractions from coking, visbreaking, steam cracking or fluid catalytic cracking (FCC) units, for example. This gasoline can optionally consist of a large proportion of gasoline originating from other production processes, such as atmospheric distillation (gasoline from direct distillation (or straight-run gasoline)) or conversion processes (coker or steam cracker gasoline). The feedstock preferably consists of a gasoline fraction produced by a fluid catalytic cracking unit.

[0038] The feedstock is a gasoline fraction containing sulfur compounds and olefins, whose boiling range generally extends from the boiling point of hydrocarbons having 2 or 3 carbon atoms (C2 or C3) to 260 °C, preferably from the boiling point of hydrocarbons having 2 or 3 carbon atoms (C2 or C3) to 220 °C, more preferably from the boiling point of hydrocarbons having 5 carbon atoms to 220 °C. The method according to the invention can also process feedstocks having an end point lower than those mentioned above, such as, for example, a C5-180 °C fraction.

[0039] The sulfur content of gasoline fractions produced by fluid catalytic cracking (FCC) depends on the sulfur content of the feedstock treated by said FCC, on the presence or absence of pretreatment of the feedstock of said FCC, and on the end point of said fraction. Generally, the sulfur content of the whole gasoline fraction, especially those originating from FCC, is greater than 100 weight ppm, and most of the time greater than 500 weight ppm. For gasoline having an end point greater than 200 °C, the sulfur content is generally greater than 1000 weight ppm; in some cases, they can even reach values of about 4000 to 5000 weight ppm.

[0040] The feedstock treated by the method according to the invention can be a feedstock containing sulfur compounds with a sulfur content greater than 200 weight ppm and often greater than 500 ppm.

[0041] In addition, gasoline produced by a fluid catalytic cracking (FCC) unit contains on average 0.5 wt% to 5 wt% of diolefins, 20 wt% to 50 wt% of olefins and 10 weight ppm to 0.5 wt% of sulfur, usually including less than 300 ppm of mercaptans.

[0042] Step a0) Selective hydrogenation (optional)

[0043] Depending on the type of gasoline to be treated, it may be advantageous to pre-treat the aforementioned gasoline in the presence of hydrogen and a selective hydrogenation catalyst, so as to at least partially hydrogenate the diolefins and to react with the olefins to increase the molecular weight of a part of the light mercaptans (RSH) present in the feedstock to obtain a reaction of thioethers.

[0044] To this end, the gasoline to be treated is sent to a selective hydrogenation catalytic reactor, which contains at least one fixed bed or moving bed of a catalyst for the selective hydrogenation of diolefins and for increasing the molecular weight of light mercaptans. The reaction for the selective hydrogenation of diolefins and for increasing the molecular weight of light mercaptans is preferably carried out on a sulfided catalyst comprising at least one Group VIII element and optionally at least one Group VIB element and an oxide support. The Group VIII element is preferably selected from nickel and cobalt, and especially nickel. The Group VIB element, when present, is preferably selected from molybdenum and tungsten, and very preferably molybdenum.

[0045] The oxide support of the catalyst is preferably selected from alumina, nickel aluminate, silica, silicon carbide or a mixture of these oxides. Alumina is preferably used, and still more preferably high-purity alumina is used. According to a preferred embodiment, the selective hydrogenation catalyst contains 1% to 12% nickel oxide (in the form of NiO) by weight of nickel and 6% to 18% molybdenum oxide (in the form of MoO3) by weight of molybdenum, with a nickel / molybdenum molar ratio of 0.3 to 2.5, and the metals are deposited on a support composed of alumina. The degree of sulfidation of the metals constituting the catalyst is preferably greater than 60%.

[0046] During the optional selective hydrogenation step, at a temperature of 50 °C to 250 °C, preferably 80 °C to 220 °C and still more preferably 90 °C to 200 °C, the gasoline is brought into contact with the catalyst at a space velocity (HSV) of 0.5 h -1 to 20 h -1 . The unit of the space velocity is the volume flow rate (l / l / h) of the raw material at 15 °C per catalyst bed volume. The pressure is 0.2 to 5 MPa, preferably 0.6 to 4 MPa and still more preferably 1 to 3 MPa. The optional selective hydrogenation step is generally carried out at a ratio of the hydrogen flow rate expressed as standard m 3 / m 3 , preferably 3 to 30 Nm 3 / m 3 to the volume flow rate of the raw material to be treated expressed as standard conditions (15 °C, 0.1 MPa) in m 3 / h. 3

[0047] After the selective hydrogenation, the diolefin content determined via the maleic anhydride value (MAV) according to the UOP 326 method is generally reduced to less than 6 mg maleic anhydride / g, actually even less than 4 mg MA / g and more preferably less than 2 mg MA / g. In some cases, less than 1 mg MA / g can be obtained.

[0048] Subsequently, the selectively hydrogenated gasoline can be distilled into at least two fractions, a light fraction, a heavy fraction and an optional middle fraction. In the case of fractionation into two fractions, the heavy fraction is treated by the process according to the invention. In the case of fractionation into three fractions, the middle fraction and the heavy fraction can be treated separately by the process according to the invention.

[0049] It should be noted that it is conceivable to simultaneously carry out the hydrogenation of the diolefins and the fractionation into two or three fractions by means of a catalytic distillation column, which comprises a distillation column equipped with at least one catalytic bed.

[0050] Step a) Selective hydrodesulfurization (HDS)

[0051] The hydrodesulfurization step a) is carried out in order to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H2S.

[0052] The temperature is generally from 200 °C to 350 °C and preferably from 220 °C to 320 °C. The temperature used must be sufficient to keep the gasoline to be treated in the gas phase in the reactor.

[0053] The operating pressure of this step is generally from 0.2 MPa to 5 MPa and preferably from 1 MPa to 3 MPa.

[0054] The amount of catalyst used in each reactor of the first reaction section is generally such that the volume flow rate ratio (expressed as m 3 of the gasoline to be treated per catalyst bed at 15 °C (expressed as m 3 / h) (also known as the space velocity or HSV) is from 1 to 20 h -1 and preferably from 2 to 10 h -1 .

[0055] The hydrogen flow rate is generally such that the ratio of the hydrogen flow rate expressed as standard m 3 / h (Nm 3 / h) to the volume flow rate of the feedstock to be treated expressed as m 3 / h under standard conditions (15 °C, 0.1 MPa) is from 10 to 1000 Nm 3 / m 3 、preferably from 50 to 600 Nm 3 / m 3 . Standard m 3 is understood to mean the volume of 1 m 3 of gas at 0 °C and 0.1 MPa.

[0056] The hydrogen required for this step can be fresh hydrogen or recycled hydrogen, preferably with H2S removed, or a mixture of fresh hydrogen and recycled hydrogen. Preferably, a mixture of fresh hydrogen and recycled hydrogen will be used.

[0057] Depending on the sulfur content of the feedstock to be treated, the degree of desulfurization in step a) is generally greater than 50% and preferably greater than 70%, such that the product resulting from step a) contains less than 200 ppm by weight of sulfur and preferably less than 100 ppm by weight of sulfur.

[0058] In the process according to the invention, the degree of hydrogenation of the olefins during this step is preferably less than 50%, more preferably less than 40%.

[0059] According to the invention, as described below, the hydrodesulfurization catalyst of step a) comprises an active phase and an oxide support, the active phase comprising at least one Group VIB metal and at least one Group VIII metal, optionally phosphorus, and preferably consisting of at least one Group VIB metal and at least one Group VIII metal, optionally phosphorus.

[0060] The Group VIB metal present in the active phase of the catalyst is preferably selected from molybdenum and tungsten.

[0061] The Group VIII metal present in the active phase of the catalyst is preferably selected from cobalt, nickel, and mixtures of these two elements.

[0062] The active phase of the catalyst is preferably selected from the group formed by the combinations of elemental nickel-molybdenum, cobalt-molybdenum, and nickel-cobalt-molybdenum, and very preferably the active phase consists of cobalt and molybdenum.

[0063] The content of the Group VIII metal, calculated as the oxide of the Group VIII metal, is preferably 0.1% to 10% by weight relative to the total weight of the catalyst, more preferably 0.6% to 8% by weight relative to the total weight of the catalyst, even more preferably 0.6% to 7% by weight, and very preferably 1% to 6% by weight. When the metal is cobalt or nickel, the metal content is expressed as CoO or NiO.

[0064] The total content of the Group VIB metal, calculated as the oxide of the Group VIB metal, is preferably 1% to 20% by weight relative to the total weight of the catalyst, more preferably 2% to 18% by weight relative to the total weight of the catalyst, and very preferably 3% to 16% by weight. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 or WO3.

[0065] Preferably, the molar ratio of the Group VIII metal to the Group VIB metal of the catalyst is generally 0.1 to 0.8 mol / mol, preferably 0.2 to 0.6 mol / mol.

[0066] Optionally, relative to the total weight of the catalyst, the catalyst may additionally have a phosphorus content of P2O5 generally from 0.3 wt% to 10 wt%, preferably from 0.3 wt% to 5 wt%, very preferably from 0.5 wt% to 3 wt%.

[0067] In addition, when phosphorus is present, the phosphorus / (Group VIB metal) molar ratio is generally from 0.1 to 0.7 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0068] Preferably, the catalyst of step a) has a specific surface area of 60 to 250 m 2 / g, preferably 60 to 200 m 2 / g and even more preferably 65 to 180 m 2 / g, and even more preferably 70 to 130 m 2 / g.

[0069] The total pore volume of the catalyst of step a) is generally from 0.3 cm 3 / g to 1.3 cm 3 / g, preferably from 0.4 cm 3 / g to 1.1 cm 3 / g.

[0070] The oxide support of the hydrodesulfurization catalyst is generally a porous solid selected from the group consisting of: alumina, silica, silica-alumina, or titanium oxide or magnesia used alone or in admixture with alumina or silica-alumina. It is preferably selected from the group consisting of silica, alumina, and silica-alumina. Very preferably, the oxide support consists essentially of alumina, that is, relative to the total weight of the support, it contains at least 51 wt%, preferably at least 60 wt%, very preferably at least 80 wt%, or even at least 90 wt% of alumina. It preferably consists only of alumina.

[0071] In a preferred embodiment, the catalyst of step a) comprises an alumina support and an active phase, the active phase comprising cobalt and molybdenum and optionally phosphorus, preferably consisting of cobalt and molybdenum and optionally phosphorus, the catalyst containing a weight content of cobalt oxide in the form of CoO of from 0.1 wt% to 10 wt%, preferably from 0.6 wt% to 8 wt%, more preferably from 0.6 wt% to 7 wt%, and even more preferably from 1 wt% to 6 wt% relative to the total weight of the catalyst, and a weight content of molybdenum oxide in the form of MoO3 of from 1 wt% to 20 wt%, preferably from 2 wt% to 18 wt%, and very preferably from 3 wt% to 16 wt% relative to the total weight of the catalyst, wherein the cobalt / molybdenum molar ratio is from 0.1 to 0.8 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0072] Preferably, the support of the hydrodesulfurization catalyst has a specific surface area of 60 to 250 m 2 / g, preferably 60 to 200 m 2 / g, and even more preferably 65 to 180 m 2 / g, and even more preferably 70 to 130 m 2 / g.

[0073] The total pore volume of the support of the hydrodesulfurization catalyst is generally 0.3 cm 3 / g to 1.3 cm 3 / g, preferably 0.4 cm 3 / g to 1.1 cm 3 / g.

[0074] The support of the hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometric shape, tablets, pellets, compressed cylinders, crushed solids or any other shaped body. Preferably, the support is in the form of beads having a diameter of 0.5 to 6 mm, or in the form of cylindrical, trilobal or quadrilobal extrudates having an outside diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0075] Subsequently, the first part of the desulfurized effluent obtained at the end of step a) is sent directly to step b) of the process according to the invention without separation.

[0076] Step b) First fine hydrodesulfurization step (FNS1)

[0077] During the hydrodesulfurization step a), most of the sulfur compounds are converted to H2S. The remaining sulfur compounds are essentially refractory sulfur compounds, as well as the re - formed mercaptans resulting from the addition of the H2S formed in step a) to the olefins present in the feedstock.

[0078] Step b) of the process according to the invention comprises converting at least a part of the re - formed mercaptans contained in the first effluent from step a) into olefins and H2S, and converting at least a part of the sulfur compounds (such as thiophene compounds) contained in the first effluent from step a) into saturated compounds, for example into tetrahydrothiophene (or thiacyclopentane) or into mercaptans, and subsequently at least partially hydrogenolyzing these sulfur compounds to form H2S.

[0079] Preferably, step b) is carried out at a higher temperature than step a). Specifically, by using a higher temperature in this step compared to the temperature of step a), the formation of mercaptans is disadvantaged by changing the thermodynamic equilibrium. Step b) also enables the continued hydrodesulfurization of the residual sulfur compounds.

[0080] The temperature is generally from 250 °C to 400 °C, preferably from 270 °C to 390 °C. The temperature employed must be sufficient to keep the gasoline to be treated in the gas phase in the reactor.

[0081] The operating pressure at this stage is generally from 0.2 MPa to 5 MPa and preferably from 1.5 MPa to 3 MPa.

[0082] The amount of catalyst used in each reactor of the second reaction section is generally such that the volume flow rate ratio of the gasoline to be treated in the catalyst bed (expressed in m 3 at standard conditions (15 °C, 0.1 MPa)) (also known as the space velocity or HSV) is from 1 to 40 h 3 / h) and preferably from 2 to 20 h -1 。 -1 .

[0083] The first hydrodesulfurization catalyst of step b) comprises an active phase composed of a Group VIII metal and an oxide support, and the active phase is at least partially sulfided.

[0084] The Group VIII metal is preferably nickel. When the Group VIII metal is nickel, the nickel sulfide phase diagram has a large number of sulfur-rich phases and nickel-rich phases at low temperatures. Thus, there may be multiple nickel sulfide phases and stoichiometries, 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, namely hexagonal α-NiS stable at high temperatures and rhombohedral β-NiS stable at low temperatures. The existence of these numerous phases makes the synthesis of nickel sulfide in single-phase form complex, and thus the product is usually a mixture of two or more phases.

[0085] The content of the Group VIII metal, calculated as the oxide of the Group VIII metal, is preferably from 5% by weight to 65% by weight relative to the total weight of the catalyst, more preferably from 8% by weight to 55% by weight relative to the total weight of the catalyst, and very preferably from 12% by weight to 40% by weight. When the metal is nickel, the metal content is expressed as NiO.

[0086] Preferably, the catalyst of step b) is characterized in that its specific surface area is from 60 to 250 m 2 / g, preferably from 70 to 200 m 2 / g.

[0087] 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 cm3 between / g.

[0088] The oxide support of the first hydrodesulfurization catalyst is generally a porous solid selected from the group consisting of: alumina, silica, silica - alumina, or titanium oxide or magnesium oxide used alone or in admixture with alumina or silica - alumina. It is preferably selected from the group consisting of silica, alumina, and silica - alumina. Very preferably, the oxide support consists essentially of alumina, that is, relative to the total weight of the support, it contains at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight of alumina. It preferably consists only of alumina.

[0089] Preferably, the support of the first hydrodesulfurization catalyst has a specific surface area of 60 to 250 m 2 / g, preferably 70 to 200 m 2 / g.

[0090] The total pore volume of the support of the first hydrodesulfurization catalyst is generally from 0.3 cm 3 / g to 1.3 cm 3 / g, preferably 0.4 cm 3 / g to 1.1 cm 3 / g.

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

[0092] Then, the second part of the desulfurized effluent obtained at the end of step b) is sent directly to step c) of the process of the present invention without separation.

[0093] Step c) Second fine hydrodesulfurization step (FNS2)

[0094] The sulfur compounds remaining after completion of step b) are substantially refractory sulfur compounds. Step c) of the process according to the present invention essentially comprises converting at least part of the refractory sulfur compounds (such as thiophene compounds) contained in the effluent from step b) into saturated compounds, for example, into tetrahydrothiophene (or thiacyclopentane) or into mercaptans, and then at least partially hydrocracking these sulfur compounds to form H2S.

[0095] The temperature is generally between 250 °C and 400 °C, preferably between 270 °C and 390 °C. The temperature employed must be sufficient to keep the gasoline to be treated in the gas phase in the reactor.

[0096] The operating pressure of this step is generally between 0.2 MPa and 5 MPa and preferably between 1.5 MPa and 3 MPa.

[0097] The amount of catalyst used in each reactor is generally such that the ratio (also known as the space velocity) expressed as the volume flow rate of gasoline to be treated per m 3 at standard conditions (15 °C, 0.1 MPa) of the catalyst bed and expressed as m 3 / h is 1 to 40 h -1 and preferably 2 to 20 h -1 .

[0098] In the process according to the invention, during this step, the total olefin hydrogenation degree of steps b) and c) is preferably less than 30%.

[0099] Depending on the sulfur content of the feedstock to be treated, the total desulfurization degree of steps b) and c) is generally greater than 50% and preferably greater than 70%, such that the product produced in step c) contains less than 50 weight ppm of sulfur and preferably less than 20 weight ppm of sulfur, and even more preferably less than 10 weight ppm of sulfur.

[0100] The second hydrodesulfurization catalyst of step c) comprises an active phase and an oxide support, the active phase comprising at least one Group VIB metal and at least one Group VIII metal, optionally phosphorus, at least partially in the form of sulfides, preferably consisting of at least one Group VIB metal and at least one Group VIII metal, optionally phosphorus, at least partially in the form of sulfides, as described below.

[0101] The Group VIB metal is preferably selected from molybdenum and tungsten. The Group VIII metal is preferably selected from cobalt, nickel and mixtures of these two elements. The active phase of the catalyst is preferably selected from the group formed by combinations of elemental nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum, and very preferably the active phase consists of cobalt and molybdenum.

[0102] The content of the Group VIII metal, calculated as the oxide of the Group VIII metal, is preferably 0.1% to 10% by weight, more preferably 0.6% to 8% by weight, even more preferably 0.6% to 7% by weight, and very preferably 1% to 6% by weight relative to the total weight of the catalyst. When the metal is cobalt or nickel, the metal content is expressed as CoO or NiO.

[0103] The total content of the Group VIB metal, calculated as the oxide of the Group VIB metal, is preferably 1% to 20% by weight, more preferably 2% to 18% by weight, and very preferably 3% to 16% by weight relative to the total weight of the catalyst. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 or WO3.

[0104] The molar ratio of the Group VIII metal to the Group VIB metal of the catalyst is usually from 0.1 to 0.8 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0105] Optionally, relative to the total weight of the catalyst, the catalyst may additionally have a phosphorus content of P2O5 which is usually from 0.3 wt% to 10 wt%, preferably from 0.3 wt% to 5 wt%, very preferably from 0.5 wt% to 3 wt%. In addition, when phosphorus is present, the phosphorus / (Group VIB metal) molar ratio is usually from 0.1 to 0.7 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0106] Preferably, the catalyst of step c) has a specific surface area of 60 to 250 m 2 / g, preferably 60 to 200 m 2 / g and even more preferably 65 to 180 m 2 / g, and even more preferably 70 to 130 m 2 / g.

[0107] The total pore volume of the catalyst of step c) is usually from 0.3 cm 3 / g to 1.3 cm 3 / g, preferably from 0.4 cm 3 / g to 1.1 cm 3 / g.

[0108] The oxide support of the hydrodesulfurization catalyst is usually a porous solid selected from the group consisting of: alumina, silica, silica-alumina, or titanium oxide or magnesium oxide used alone or in combination with alumina or silica-alumina. It is preferably selected from the group consisting of silica, alumina, and silica-alumina. Very preferably, the oxide support consists essentially of alumina, that is, relative to the total weight of the support, it contains at least 51 wt%, preferably at least 60 wt%, very preferably at least 80 wt%, or even at least 90 wt% of alumina. It preferably consists only of alumina.

[0109] In a preferred embodiment, the catalyst of step c) comprises an alumina support and an active phase, the active phase comprising cobalt and molybdenum, preferably consisting of cobalt and molybdenum, the catalyst containing a weight content of cobalt oxide in the form of CoO of from 0.1% to 10% by weight, preferably from 0.6% to 8% by weight, more preferably from 0.6% to 7% by weight and even more preferably from 1% to 6% by weight relative to the total weight of the catalyst, and a weight content of molybdenum oxide in the form of MoO3 of from 1% to 20% by weight, preferably from 2% to 18% by weight, and very preferably from 3% to 16% by weight relative to the total weight of the catalyst, wherein the cobalt / molybdenum molar ratio is from 0.1 to 0.8 mol / mol, preferably from 0.2 to 0.6 mol / mol.

[0110] Preferably, the support of the second hydrodesulfurization catalyst has a specific surface area of 60 to 250 m 2 / g, preferably 60 to 200 m 2 / g, and even more preferably 65 to 180 m 2 / g, and even more preferably 70 to 130 m 2 / g.

[0111] The total pore volume of the support of the second hydrodesulfurization catalyst is generally from 0.3 cm 3 / g to 1.3 cm 3 / g, preferably from 0.4 cm 3 / g to 1.1 cm 3 / g.

[0112] The support of the second hydrodesulfurization catalyst may be in the form of beads, extrudates of any geometric shape, pellets, granules, compressed cylinders, crushed solids or any other shaped body. Preferably, the support is in the form of beads having a diameter of 0.5 to 6 mm, or in the form of cylindrical, trilobal or quadrilobal extrudates having an outside diameter of 0.8 to 3 mm. More preferably, the support is in the form of beads.

[0113] In a preferred scenario, the catalyst used in step c) is the same as the catalyst used in step a).

[0114] Preferably, step c) is carried out in the same reactor as step b).

[0115] Steps b) and c) for implementing the method

[0116] Steps b) and c) of the process according to the invention may be carried out in one, two or more reactors.

[0117] 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 hydrofining hydrodesulfurization reactor which contains a second reaction section through which a partial desulfurized effluent produced in step a) passes, and then step c) can be carried out in a second hydrofining hydrodesulfurization reactor which contains a third reaction section located downstream of said first reactor.

[0118] 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 which contains a second reaction section, and step c) is carried out in a second zone which contains a third reaction section downstream of the first zone.

[0119] According to one or more embodiments, the second reaction section containing the first hydrofining hydrodesulfurization catalyst occupies a volume V1, and the third hydrofining hydrodesulfurization reaction section containing the second hydrofining hydrodesulfurization catalyst occupies a volume V2, and the distribution of the volumes V1 / V2 of the second and third hydrofining hydrodesulfurization reaction sections respectively is from 90% by volume / 10% by volume to 10% by volume / 90% by volume, preferably from 90% by volume / 10% by volume to 60% by volume / 40% by volume.

[0120] Step d): Separation of H2S [optional]

[0121] This separation step d) serves to separate excess hydrogen and H2S formed during steps a), b) and c). Any method known to those skilled in the art can be envisaged.

[0122] According to a first embodiment, after steps a), b) and c), the third desulfurized effluent produced in step c) is cooled to a temperature generally below 80 °C to condense hydrocarbons. The gas phase and the liquid phase are then separated in a separation drum. The liquid fraction containing desulfurized gasoline and a part of the dissolved H2S is sent to a stabilizer column or a debutanizer column. This column separates a top fraction consisting essentially of residual H2S and hydrocarbon compounds having a boiling point lower than or equal to the boiling point of butane and a bottom fraction called stabilized gasoline containing compounds having a boiling point higher than the boiling point of normal butane, from which H2S has been removed.

[0123] According to a second embodiment, after the condensation step, the liquid fraction containing desulfurized gasoline and a part of the dissolved H2S is sent to a stripping section, while the gas fraction consisting mainly of hydrogen and H2S is sent to a purification section. Stripping can be carried out by heating said hydrocarbon fraction alone or in the presence of injected hydrogen or steam in a distillation column in order to extract at the top the light compounds entrained by dissolution in the liquid fraction and the dissolved residual H2S. The temperature of the stripped gasoline recovered at the bottom is generally from 120 °C to 250 °C.

[0124] Preferably, the separation step d) is carried out in a stabilizer column or a debutanizer column. This is because the stabilizer column enables more efficient separation of H2S than the stripping section.

[0125] Step d) is preferably carried out such that the sulfur in the form of H2S remaining in the desulfurized gasoline accounts for less than 30%, preferably less than 20%, and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.

[0126] Preparation of the catalyst

[0127] The catalyst used in the process according to the invention can be prepared by any technique known to those skilled in the art, in particular by impregnating a selected porous support with elements of Group VIII and optionally elements of Group VIB and phosphorus. The impregnation can be carried out, for example, according to the method known to those skilled in the art under the term dry impregnation, in which exactly the required precursor amounts of the elements in the form of salts soluble in the selected solvent (e.g., softened water) are introduced, so as to fill the pores of the support as precisely as possible. Preferably, when the impregnating aqueous solution contains cobalt, molybdenum, and phosphorus, the impregnating aqueous solution is prepared under pH conditions that promote the formation of dissolved heteropolyanions. For example, the pH of such an aqueous solution is from 1 to 5. Preferably, the preparation of the catalyst is carried out without mixing and adding organic reagents with the precursors of the elements of Group VIII, Group VIB, and phosphorus.

[0128] For example, in the molybdenum source, oxides and hydroxides, molybdic acid and its salts, especially ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PMo 12 O 40 ) and its salts, and optionally silicomolybdic acid (H4SiMo 12 O 40 ) and its salts can be used. The molybdenum source can also be, for example, any heteropoly compound of the Keggin, lacunary Keggin, substituted Keggin, Dawson, Anderson, or Strandberg type. Preferably, molybdenum trioxide and heteropoly compounds of the Keggin, lacunary Keggin, substituted Keggin, and Strandberg types are used.

[0129] The cobalt precursors that can be used are advantageously selected from, for example, oxides, hydroxides, hydroxycarbonates, carbonates, and nitrates. Cobalt hydroxide and cobalt carbonate are preferably used.

[0130] The nickel precursors that can be used are advantageously selected from, for example, oxides, hydroxides, hydroxycarbonates, carbonates, and nitrates. Nickel hydroxide and nickel hydroxycarbonate are preferably used.

[0131] The tungsten precursors that can be used are also known to those skilled in the art. For example, among the tungsten sources, oxides and hydroxides, tungstic acid and its salts, in particular ammonium salts, such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and its salts, and optional silicotungstic acid (H4SiW 12 O 40 ) and their salts. The tungsten source may also be any heteropoly compound of the Keggin, vacancy Keggin, substituted Keggin or Dawson type. Oxides and ammonium salts, such as ammonium metatungstate, or heteropoly anions of Keggin, vacancy Keggin or substituted Keggin are preferably used.

[0132] Phosphorus can advantageously be introduced alone or mixed with at least one of the elements of Group VIB and Group VIII. Phosphorus is preferably introduced mixed with precursors of the elements of Group VIB and Group VIII by dry impregnation of the porous support with a solution containing the element precursors and the phosphorus precursors. The preferred phosphorus source is orthophosphoric acid H3PO4, but its salts and esters, such as ammonium phosphate or mixtures thereof are also suitable. Phosphorus can also be introduced simultaneously with the elements of Group VIB in the form of heteropolyanions of the Keggin, vacancy Keggin, substitution Keggin or Strandberg type.

[0133] The support thus filled with the solution may be aged at a temperature below 50° C., preferably at ambient temperature, for a period not exceeding 12 h, preferably not exceeding 6 h.

[0134] After the maturation step, the catalyst precursor obtained can be subjected to a thermal treatment. The purpose of this treatment is generally to convert the molecular precursor of the element into the oxide phase. In this case, this is an oxidation treatment, but a simple drying of the catalyst can also be carried out.

[0135] In case of drying, the catalyst precursor is dried at a temperature of 50°C to below 200°C, preferably 70°C to 180°C, for a period of time of generally 0.5h to 12h, and even more preferably 0.5h to 5h.

[0136] In the case of oxidation treatment (also called calcination), the treatment is generally carried out under air or under diluted oxygen, and the treatment temperature is generally 200°C to 550°C, preferably 300°C to 500°C, and advantageously lasts for a period of generally 0.5h to 24h, preferably for a period of 0.5h to 12h, and even more preferably for a period of 0.5h to 10h.

[0137] Before it is used as a hydrotreating catalyst, it is advantageous to subject the optionally dried or calcined catalyst to a step of activation by sulfidation. This activation stage is carried out by methods known to those skilled in the art and is advantageously carried out in a sulfur-reducing atmosphere in the presence of hydrogen and hydrogen sulfide. Hydrogen sulfide can be used directly or generated from a sulfide reagent such as dimethyldisulfide.

[0138] Description of the sulfidation of the catalyst

[0139] Before contacting the feedstock to be treated in a hydrodesulfurization process for gasoline, the catalyst used in steps a0), a), b) and c) of the process according to the invention generally undergoes a sulfidation step. The sulfidation is preferably carried out in a sulfur-reducing medium, that is to say in the presence of H2S and hydrogen, in order to convert the metal oxides into sulfides such as for example MoS2, Co9S8 or Ni3S2. The sulfidation is carried out by injecting into the catalyst a stream containing H2S and hydrogen, or a sulfur compound and hydrogen capable of decomposing into H2S in the presence of the catalyst. Polysulfides, such as dimethyldisulfide (DMDS), are H2S precursors commonly used to sulfide the catalysts of steps a0), a), b) and c). The sulfur can also originate from the feedstock. The temperature is adjusted so that H2S reacts with the 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 a temperature of 200 °C to 600 °C and more preferably 300 °C to 500 °C.

[0140] The degree of sulfidation of the metals constituting the catalysts of steps a0), a), b) and c) is at least equal to 60%, preferably at least equal to 70%. The sulfur content in the catalysts of the sulfidation steps a0), a), b) and c) is measured by elemental analysis according to ASTM D5373. A metal is considered to be sulfided when the overall degree of sulfidation defined by the molar ratio of sulfur (S) present on the catalyst to the metal is at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of the metal under consideration. The total degree of sulfidation is defined by the following equation:

[0141] (S / metal) 催化剂 ≥0.6 × (S / metal) 理论

[0142] where:

[0143] (S / metal) 催化剂 is the molar ratio of sulfur (S) present on the catalyst to the metal

[0144] (S / metal) 理论 is the molar ratio of sulfur to the metal corresponding to the complete sulfidation of the metal to obtain a sulfide.

[0145] The theoretical molar ratio varies according to the relevant metal:

[0146] -(S / Co) 理论 = 1

[0147] -(S / Ni) 理论 = 1

[0148] -(S / Mo) 理论 = 2 / 1

[0149] -(S / W) 理论 = 2 / 1

[0150] When the catalyst used in steps a0), a), b) and c) contains multiple metals, the molar ratio of sulfur to the combined metals present on the catalyst must also be at least equal to 60% of the theoretical molar ratio corresponding to the sulfide obtained by complete sulfidation of each metal, and this calculation is carried out in proportion to the relative molar fractions of each metal.

[0151] The following examples illustrate the invention without limiting its scope.

[0152] Examples

[0153] The analytical methods used to characterize the feedstock and the effluent are as follows:

[0154] - For sulfur contents above 10 ppm S, the sulfur content according to the ASTM D2622 method, and for sulfur contents below 10 ppm S, the sulfur content according to the ISO 20846 method;

[0155] - The mercaptan content according to the ASTM D3227 method;

[0156] - The olefin content based on gas chromatography analysis according to the ASTM D6733 method.

[0157] Example 1: Preparation of catalyst A

[0158] There is provided a support A' composed of alumina in the form of beads, having a particle size of 2 to 4 mm and having a specific surface area of 139 m 2 / g and a pore volume of 0.97 ml / g.

[0159] Subsequently, cobalt and molybdenum are added. The impregnation solution is prepared by dissolving ammonium heptamolybdate tetrahydrate (5.64 g, ≥99.5%, Sigma- ) and cobalt nitrate hexahydrate (5.36 g, ≥99.5%, Alfa ) It is prepared by. After dry-impregnating 40 g of support A', the impregnated alumina is cured in a water-saturated atmosphere at ambient temperature for 4 h, then dried at 120 °C for 4 h, and finally calcined at 450 °C for 4 h under an air flow of 1 l / h / g. The catalyst thus obtained is designated as A.

[0160] Expressed in oxide form and relative to the weight of the dry catalyst, the final elemental composition of catalyst A is thus as follows: MoO3 = 10.0 ± 0.2 wt% and CoO = 3.0 ± 0.1 wt%.

[0161] The Co / Mo and P / Mo molar ratios are 0.60 and 0, respectively.

[0162] The specific surface area of catalyst A is 124 m 2 / g.

[0163] Example 2: Preparation of catalyst B

[0164] Support B' identical to support A' is provided.

[0165] Nickel is then added. The impregnation solution is prepared by dissolving nickel nitrate hexahydrate (34.36 g, ≥99.5%, Sigma- ) in 25 ml of softened water at room temperature. After dry-impregnating 40 g of support B', the impregnated alumina is cured in a water-saturated atmosphere at ambient temperature for 4 h, then dried at 120 °C for 4 h, and finally calcined at 450 °C for 4 h under an air flow of 1 l / h / g. The catalyst thus obtained is designated as B.

[0166] Expressed in oxide form and relative to the weight of the dry catalyst, the final elemental composition of catalyst B is thus as follows: NiO = 17.9 ± 0.3 wt%.

[0167] The specific surface area of catalyst B is 114 m 2 / g.

[0168] Example 3: Use of the catalyst in the gasoline desulfurization process

[0169] Example 3 aims to demonstrate the advantages of a gasoline desulfurization process employing a series of steps and using a specific catalyst in each step. Gasoline from a fluid catalytic cracking unit consisting of 25 wt% olefins and 600 ppm S total sulfur is subjected to several steps of treatment:

[0170] - A step of selective hydrodesulfurization (HDS) using catalyst A in an adiabatic reactor. The operating conditions for the one-step hydrodesulfurization step of the gasoline feedstock are as follows: HSV = 3 h -1 , P = 2.0 MPa. A pure hydrogen stream is added to the feedstock at the reactor inlet such that H2 / HC = 250 Nm 3 / m 3 Direct the effluent directly to the reactor of the second step;

[0171] - The first hydrodesulfurization (FNS1) step using catalyst A or B in an adiabatic reactor. Only the effluent from the first step is processed in this second step. The pressure of the first hydrodesulfurization step is fixed at P = 2.0 MPa. The reactor inlet temperature is always fixed 35 °C higher than the temperature of the first effluent leaving the selective hydrodesulfurization step;

[0172] - Optionally, a second hydrodesulfurization (FNS2) step using catalyst A or B in an adiabatic reactor. Only the effluent from the previous step is processed in this second hydrodesulfurization step. The pressure of the second hydrodesulfurization step is fixed at P = 2.0 MPa. The reactor inlet temperature is equal to the temperature of the effluent leaving the previous step.

[0173] Fix the inlet temperature of the selective hydrodesulfurization reactor to obtain an effluent containing 10 weight ppm S of total sulfur (i.e., a total sulfur conversion rate of greater than 98%). Before use, the catalysts contained in the selective and hydrodesulfurization reactors are sulfided by contacting with a feedstock consisting of sulfur in the form of 2 wt% dimethyldisulfide (DMDS) in n-heptane at a pressure of 3.4 MPa and a temperature of 350 °C for 4 h.

[0174] The results show that, compared with the prior art, the gasoline hydrodesulfurization method of the present invention can achieve the best performance because the embodiments of the present invention can maximize the reduction of the increase in the average temperature of the HDS section while increasing the olefin content at the outlet of the method, thereby extending the service life of the catalyst. The performance characteristics of the gasoline hydrodesulfurization method are listed in Table 1.

[0175] Table 1

[0176]

[0177]

[0178] The "average temperature" of the HDS, FNS1, and FNS2 steps corresponds to the weighted average bed temperature (WABT) well known to those skilled in the art. The average temperature is advantageously determined as a function of the catalytic system used, the equipment items, and their configuration. The average temperature (or WABT) is calculated as follows:

[0179] WABT = (T 入口 + T 出口 ) / 2

[0180] where T 入口 : the temperature at the inlet of the reaction section, T 出口: The temperature at the outlet of the reaction section. Unless otherwise stated, the "average temperature" of the reaction section is given under recycle startup conditions.

Claims

1. A method for treating gasoline containing sulfur compounds and olefins, said method comprising at least the following steps: a) In the first reaction stage, gasoline, hydrogen, and a hydrodesulfurization catalyst are contacted at a temperature of 200°C to 350°C and a pressure of 0.2 MPa to 5 MPa, with a space velocity of 1 h -1 to 20 h -1 , expressed as the ratio of the hydrogen flow rate in standard m 3 / h to the flow rate of the feedstock to be treated in standard conditions, expressed as m 3 / h, is 10 Nm 3 / m 3 to 1000 Nm 3 / m 3 , and the hydrodesulfurization catalyst comprises an active phase and an oxide support, and the active phase comprises a Group VIB metal and a Group VIII metal at least partially in the form of sulfides, thereby obtaining a first partially desulfurized effluent; b) Without separating the H2S formed in step a), in the second reaction section, the first part of the desulfurized effluent obtained at the end of step a) is directly contacted with a first hydrodesulfurization catalyst for refining at a temperature of 250 °C to 400 °C and a pressure of 0.2 MPa to 5 MPa, where the space velocity is 1 h -1 to 40 h -1 , and the first hydrodesulfurization catalyst for refining comprises an active phase and an oxide support, and the active phase is composed of Group VIII metals at least partially in the form of sulfides, thereby obtaining a second part of the desulfurized effluent; c) Without separating the H2S formed in step b), in the third reaction section, the second part of the desulfurized effluent obtained at the end of step b) is brought into direct contact with a second hydrodesulfurization catalyst for refining at a temperature of 250 °C to 400 °C and a pressure of 0.2 MPa to 5 MPa, with a space velocity of 1 h -1 to 40 h -1 , and the second hydrodesulfurization catalyst for refining comprises an active phase and an oxide support, and the active phase comprises at least one Group VIB metal and at least one Group VIII metal, at least part of which is in the form of sulfide, so as to obtain a third part of the desulfurized effluent.

2. The method according to claim 1, wherein the second reaction section containing the first hydrodesulfurization catalyst occupies a volume V1, and the third hydrodesulfurization reaction section containing the second hydrodesulfurization catalyst occupies a volume V2, and the volume distribution of V1 / V2 of the second and third reaction sections is 90% by volume / 10% by volume to 10% by volume / 90% by volume.

3. The method according to claim 1 or 2, wherein the catalyst in step a) and / or step c) contains a Group VIII metal content of 0.1% by weight to 10% by weight based on the total weight of the catalyst in terms of the oxide of the Group VIII metal, and a Group VIB metal content of 1% by weight to 20% by weight based on the total weight of the catalyst in terms of the oxide of the Group VIB metal.

4. The method according to any one of the preceding claims, wherein the catalyst in step a) and / or step c) contains alumina and an active phase containing cobalt and molybdenum, and the catalyst contains a weight content of cobalt oxide in the form of CoO of 0.1% by weight to 10% by weight based on the total weight of the catalyst, and a weight content of molybdenum oxide in the form of MoO3 of 1% by weight to 20% by weight based on the total weight of the catalyst, wherein the cobalt / molybdenum molar ratio is 0.1 to 0.8 mol / mol.

5. The method according to any one of the preceding claims, wherein the catalyst in step a) and / or step c) further contains phosphorus, and the catalyst contains a phosphorus oxide in the form of P2O5 of 0.3% by weight to 10% by weight based on the total weight of the catalyst.

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

7. The method according to any one of the preceding claims, wherein the catalysts in step a) and c) are the same.

8. The method according to any one of the preceding claims, wherein the catalyst in step b) contains a Group VIII metal content of 5% by weight to 65% by weight based on the total weight of the catalyst in terms of the oxide of the Group VIII metal.

9. The method according to any one of the preceding claims, wherein the catalyst in step b) contains an alumina support and an active phase composed of nickel, and the catalyst contains a nickel oxide in the form of NiO of 5% by weight to 65% by weight based on the total weight of the catalyst.

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

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

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

13. The method according to any one of the preceding claims, wherein the temperatures of steps b) and c) are higher than the temperature of step a).

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

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