PROCESS FOR FINIISHING HYDRODESULFURIZATION OF GASOLINES USING A CHAIN OF CATALYSTS

Abstract

A process for treating a gasoline containing sulfur compounds and olefins, comprising the following steps: a) the gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a group VIB metal and a group VIII metal, and a support, are brought into contact in a first reaction section in order to obtain a first partially desulfurized effluent; b) the first partially desulfurized effluent and a first polishing hydrodesulfurization catalyst comprising an active phase of a group VIII metal, and a support, are brought into contact in order to obtain a second partially desulfurized effluent; c) the second partially desulfurized effluent and a second polishing hydrodesulfurization catalyst comprising an active phase comprising a group VIB metal and a group VIII metal, and a support, are brought into contact in a third reaction section in order to obtain a third desulfurized effluent.

Description

FIELD OF THE INVENTION

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

PRIOR ART

Automotive fuel specifications require a significant reduction in the sulfur content in these fuels, and notably in gasolines. This reduction is notably directed toward limiting the content of sulfur and nitrogen oxides in motor vehicle exhaust gases. The specifications currently in force in Europe since 2009 for gasoline fuels set a maximum content of 10 ppm (parts per million) by weight of sulfur. Such specifications are also in force in other countries, for instance the United States and China, where the same maximum sulfur content has been required since January 2017. To achieve these specifications, it is necessary to treat gasolines via desulfurization processes.

The main sources of sulfur in gasoline bases are “cracked” gasolines, and mainly the gasoline fraction obtained from a process of catalytic cracking of a vacuum distillate or of an atmospheric or vacuum distillation residue of a crude oil. The gasoline fraction from catalytic cracking, which represents on average 40% of gasoline bases, in fact accounts for more than 90% of the sulfur in gasolines. Consequently, the production of low-sulfur gasolines requires a step of desulfurization of the catalytic cracking gasolines. Among the other sources of gasolines that may contain sulfur, mention may also be made of coker gasolines, visbreaker gasolines or, to a lesser extent, gasolines obtained from atmospheric distillation or steam cracking gasolines.

The removal of sulfur from gasoline cuts consists in specifically treating these sulfur-rich gasolines via desulfurization processes in the presence of hydrogen. These are then referred to as hydrodesulfurization (HDS) processes. However, these gasoline cuts, and more particularly the fluid catalytic cracking (FCC) gasolines, contain a large proportion of unsaturated compounds in the form of monoolefins (about 20% to 50% by weight) which contribute toward a good octane number, diolefins (0.5% to 5% by weight) and aromatics.

These unsaturated compounds are unstable and react during the hydrodesulfurization treatment. Diolefins form gums by polymerization during the hydrodesulfurization treatments. This gum formation leads to gradual deactivation of the hydrodesulfurization catalysts or gradual clogging of the reactor. Consequently, the diolefins must be removed by hydrogenation before any treatment of these gasolines. Conventional treatment processes desulfurize gasolines non-selectively by hydrogenating a large portion of the monoolefins, giving rise to a high loss of octane number and high hydrogen consumption. The most recent hydrodesulfurization processes make it possible to desulfurize cracked gasolines rich in monoolefins, while at the same time 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.

However, when very thorough desulfurization of cracked gasolines needs to be performed, some of the olefins present in the cracked gasolines are hydrogenated, on the one hand, and recombine with H₂S to form mercaptans, on the other hand. This family of compounds, of chemical formula R—SH where R is an alkyl group, are generally called recombinant mercaptans, and generally represent between 20% by weight and 80% by weight of the residual sulfur in desulfurized gasolines. Reduction of the content of recombinant mercaptans may be achieved by catalytic hydrodesulfurization, but this leads to the hydrogenation of a large portion of the monoolefins present in the gasoline, which then leads to a large reduction in the octane number of the gasoline and also to an overconsumption of hydrogen. It is moreover known that the loss of octane due to the hydrogenation of the monoolefins during the hydrodesulfurization step is proportionately greater the lower the targeted sulfur content, i.e. when it is sought to thoroughly remove the sulfur compounds present in the feedstock.

It is thus possible to treat the gasoline by a sequence of two reactors as described in document EP 1 077 247; the aim of the first stage, also called the selective HDS stage, is generally to carry out a deep desulfurization of the gasoline with minimal olefin saturation (and no aromatic loss), resulting in a maximum octane retention. The catalyst employed is generally a catalyst of CoMo type. During this stage, new sulfur compounds are formed by recombination of the H₂S resulting from the desulfurization and the olefins: recombinant mercaptans.

The second stage generally has the role of minimizing the amount of recombinant mercaptans.

The temperature is generally higher in the second stage 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 to a temperature greater than that of the first.

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

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

Surprisingly, the applicant has identified that a sequence of two specific catalysts of different nature and in a particular order in the polishing hydrodesulfurization section located downstream of the selective hydrodesulfurization (HDS) section has a synergistic effect in terms of selectivity while minimizing the saturation of the olefins which leads to a loss in octane number. Specifically, the choice of the right active phase and a suitable support makes it possible to observe a synergy between a first catalyst in the polishing section consisting of an active phase based on a group VIII element, enabling the recombinant mercaptans to be removed while preserving the olefins and a second catalyst in the polishing section consisting of an active phase of a group VIII element and a group VIB element, enabling the more refractory sulfur compounds to be removed. Moreover, the introduction of a second catalyst into the polishing section consisting of an active phase of a group VIII element and a group VIB element makes it possible to reduce the average treatment temperature in the HDS section and thus to increase the overall cycle time of the process.

SUBJECT MATTERS OF THE INVENTION

The aim of the present invention is to implement a process for producing gasolines having a low content of sulfur, making it possible to upgrade the entirety of a sulfur-containing gasoline cut, preferably a catalytic cracking gasoline cut, and to reduce the contents of sulfur in said gasoline cut to very low levels without reducing the gasoline yield and while minimizing the reduction in the octane number caused by hydrogenation of the olefins.

Thus, a subject of the present invention is a process for treating a gasoline containing sulfur compounds and olefins, the process comprising at least the following steps:

• • a) the gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a group VIB metal and a group Vill metal at least partly in sulfide form, and an oxide support, are brought into contact in a first reaction section at a temperature of between 200° C. and 350° C., at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 h⁻¹ and 20 h⁻¹ and a ratio of the hydrogen flow rate, expressed in normal m³ per hour, to the flow rate of feedstock to be treated, expressed in m³ per hour under standard conditions, of between 10 Nm³/m³ and 1000 Nm³/m³, in order to obtain a first partially desulfurized effluent;
  • b) without separating the H₂S formed in step a), the first partially desulfurized effluent obtained on conclusion of step a) and a first polishing hydrodesulfurization catalyst comprising an active phase consisting of a group Vill metal at least partly in sulfide form, and an oxide support, are brought directly into contact in a second reaction section at a temperature of between 250° C. and 400° C., at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 h⁻¹ and 40 h⁻¹, in order to obtain a second partially desulfurized effluent;
  • c) without separating the H₂S formed in step b), the second partially desulfurized effluent obtained on conclusion of step b) and a second polishing hydrodesulfurization catalyst comprising an active phase comprising, preferably consisting of, at least one group VIB metal and at least one group VIII metal at least partly in sulfide form, and an oxide support, are brought directly into contact in a third reaction section at a temperature of between 250° C. and 400° C., at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 h⁻¹ and 40 h⁻¹, in order to obtain a third partially desulfurized effluent.

According to one or more embodiments, said second reaction section containing the first polishing hydrodesulfurization catalyst occupies a volume V1, and said third polishing hydrodesulfurization reaction section containing the second polishing hydrodesulfurization catalyst occupies a volume V2, the distribution of the 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 polishing hydrodesulfurization reaction section.

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

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 content by weight, relative to the total weight of catalyst, of cobalt oxide, in CoO form, of between 0.1% and 10%, a content by weight, relative to the total weight of catalyst, of molybdenum oxide, in MoO₃ form, of between 1% and 20%, and a cobalt/molybdenum molar ratio of between 0.1 and 0.8 mol/mol.

According to one or more embodiments, the catalyst of step a) and/or of step c) further comprises phosphorus, said catalyst contains a content by weight, relative to the total weight of the catalyst, of phosphorus oxide, in P₂O₅ form, of between 0.3% and 10% by weight.

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

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

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

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

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

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

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

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

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

DETAILED DESCRIPTION OF THE INVENTION

Definitions

In the text hereinbelow, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, editor-in-chief D. R. Lide, 81st edition, 2000-2001). For example, group VIII (or VIIIB) according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification.

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

In the following description of the invention, the “total pore volume” of the oxide support or of the 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 dyne/cm and a contact angle of 140°. The wetting angle was taken equal to 140° following the recommendations of the publication “Techniques de l'ingénieur, traité analyse et caractérisation” [Techniques of the Engineer, Analysis and Characterization Treatise], pages 1050-5, written by Jean Charpin and Bernard Rasneur.

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

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

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

The Feedstock

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

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

The sulfur content of the gasoline cuts produced by catalytic cracking (FCC) depends on the sulfur content of the feedstock treated by the FCC, on the presence or absence of a pretreatment of the feedstock of the FCC, as well as on the end point of the cut. Generally, the sulfur contents of the whole of a gasoline cut, in particular those originating from the FCC, are greater than 100 ppm by weight and most of the time greater than 500 ppm by weight. For gasolines having end points of greater than 200° C., the sulfur contents are often greater than 1000 ppm by weight; they can even, in certain cases, reach values of the order of 4000 to 5000 ppm by weight.

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

In addition, the gasolines resulting from catalytic cracking (FCC) units contain, on average, between 0.5% and 5% by weight of diolefins, between 20% and 50% by weight of olefins and between 10 ppm and 0.5% by weight of sulfur, including generally less than 300 ppm of mercaptans.

Step a0) Selective Hydrogenation (Optional)

Depending on the type of gasoline to be treated, it may be advantageous to treat the gasoline beforehand in the presence of hydrogen and of a selective hydrogenation catalyst so as to at least partially hydrogenate the diolefins and to carry out a reaction for increasing the molecular weight of a portion of the light mercaptans (RSH) present in the feedstock to give thioethers, by reaction with olefins.

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 catalyst for the selective hydrogenation of the diolefins and for increasing the molecular weight of the light mercaptans. The reaction for the selective hydrogenation of the diolefins and for increasing the molecular weight of the light mercaptans is preferentially 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 chosen from nickel and cobalt and in particular nickel. The group VIB element, when it is present, is preferably chosen from molybdenum and tungsten and very preferably molybdenum.

The oxide support of the catalyst is preferably chosen from alumina, nickel aluminate, silica, silicon carbide or a mixture of these oxides. Use is preferably made of alumina and more preferably still of high-purity alumina. According to a preferred embodiment, the selective hydrogenation catalyst contains nickel at a content by weight of nickel oxide, in NiO form, of between 1% and 12%, and molybdenum at a content by weight of molybdenum oxide, in MoO₃ form, of between 6% and 18% and a nickel/molybdenum molar ratio of between 0.3 and 2.5, the metals being deposited on a support consisting of alumina. The degree of sulfidation of the metals constituting the catalyst is preferably greater than 60%.

During the optional selective hydrogenation step, the gasoline is brought into contact with the catalyst at a temperature of between 50° C. and 250° C., preferably between 80° C. and 220° C. and more preferably still between 90° C. and 200° C., with an hourly space velocity (HSV) of between 0.5 h⁻¹ and 20 h⁻¹, the unit of the hourly space velocity being the volume flow rate of feedstock at 15° C. per volume of catalyst bed (l/I/h). The pressure is between 0.2 and 5 MPa, preferably between 0.6 and 4 MPa and more preferably still between 1 and 3 MPa. The optional selective hydrogenation step is typically carried out with a ratio of the hydrogen flow rate, expressed in normal m³ per hour, to the volume flow rate of feedstock 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.

After selective hydrogenation, the content of diolefins, determined via the maleic anhydride value (MAV), according to the UOP 326 method, is generally reduced to less than 6 mg maleic anhydride/g, indeed 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 may be obtained.

The selectively hydrogenated gasoline may then subsequently be distilled into at least two cuts, a light cut and a heavy cut and optionally an intermediate cut. In the case of the fractionation into two cuts, the heavy cut is treated according to the process of the invention. In the case of the fractionation into three cuts, the intermediate and heavy cuts can be treated separately by the process according to the invention.

It should be noted that it is possible to envisage carrying out the steps of hydrogenation of the diolefins and of fractionation into two or three cuts simultaneously by means of a catalytic distillation column which includes a distillation column equipped with at least one catalytic bed.

Step a) Selective Hydrodesulfurization (HDS)

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

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

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

The amount of catalyst employed in each reactor of the first reaction section is generally such that the ratio of the volume flow rate at 15° C. of gasoline to be treated, expressed in m³ per hour, per m³ of catalyst bed (also called hourly space velocity or HSV) is between 1 and 20 h⁻¹ and preferably between 2 and 10 h⁻¹.

The hydrogen flow rate is generally such that the ratio of the hydrogen flow rate, expressed in normal m³ per hour (Nm³/h), to the volume flow rate of feedstock 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 is understood to mean the volume of 1 m³ of gas at 0° C. and 0.1 MPa.

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

The degree of desulfurization of step a), which depends on the sulfur content of the feedstock to be treated, is generally greater than 50% and preferably greater than 70%, so that the product resulting from step a) contains less than 200 ppm by weight of sulfur and preferably less than 100 ppm by weight of sulfur.

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

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

The group VIB metal present in the active phase of the catalyst is preferentially chosen from molybdenum and tungsten.

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

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

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

The total content of group VIB metal is preferably between 1% and 20% by weight of oxide of the group VIB metal relative to the total weight of the catalyst, more preferentially between 2% and 18% by weight, and very preferably between 3% and 16% by weight of 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.

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

Optionally, the catalyst can additionally have a phosphorus content generally of between 0.3% and 10% by weight of P₂O₅ relative to the total weight of catalyst, preferably between 0.3% and 5% by weight, very preferably between 0.5% and 3% by weight.

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

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

The total pore volume of the catalyst of step a) is generally between 0.3 cm³/g and 1.3 cm³/g, preferably of between 0.4 cm³/g and 1.1 cm³/g.

The oxide support of the hydrodesulfurization catalyst is typically a porous solid chosen from the group consisting of: alumina, silica, silica-alumina or else titanium or magnesium oxides, used alone or as a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, alumina and silica-alumina. Very preferably, the oxide support essentially consists of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight, of alumina, relative to the total weight of said support. It preferably consists solely of alumina.

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

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

The total pore volume of the support of the hydrodesulfurization catalyst is generally between 0.3 cm³/g and 1.3 cm³/g, preferably between 0.4 cm³/g and 1.1 cm³/g.

The support of the hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometry, platelets, of pellets, a compressed cylinder, crushed solids or any other shaping.

Preferably, the support is in the form of beads with a diameter of 0.5 to 6 mm or in the form of cylindrical, trilobe or quadrilobe extrudates with a circumscribed diameter of 0.8 to 3 mm. More preferentially, the support is in the form of beads.

The first partially desulfurized effluent obtained on conclusion of step a) is then sent directly and without separation to step b) of the process according to the invention.

Step b) First Polishing Hydrodesulfurization Step (FNS1)

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

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

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

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

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

The amount of catalyst employed in each reactor of the second reaction section is generally such that the ratio of the volume flow rate of gasoline to be treated, expressed in m³ per hour under standard conditions (15° C., 0.1 MPa), per m³ of catalyst bed (also called hourly space velocity or HSV) is between 1 and 40 h⁻¹ and preferably between 2 and 20 h⁻¹.

The first hydrodesulfurization catalyst of step b) comprises an active phase consisting of a group VIII metal, and of an oxide support, said active phase being at least partially sulfided.

The group Vill metal is preferably nickel. When the group VIII metal is nickel, the nickel sulfide phase diagram has a large number of sulfur-rich and nickel-rich phases at low temperature. Various nickel sulfide phases and stoichiometries are therefore possible, ranging from nickel-rich compounds such as Ni₃S₂, Ni₆S₅, Ni₇S₆, NipS₈ and NiS to sulfur-rich compounds like Ni₃S₄ and NiS₂. It should be noted that NiS is also known to exist in two main phases, namely the hexagonal α-NiS, which is stable at high temperatures, and the rhombohedral β-NiS, which is stable at low temperature. The existence of these numerous phases makes the synthesis of nickel sulfide in the form of a single phase complex, the products therefore often being mixtures of two or more phases.

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

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

The total pore volume of the catalyst of 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.

The oxide support of the first polishing hydrodesulfurization catalyst is typically a porous solid chosen from the group consisting of: alumina, silica, silica-alumina or else titanium or magnesium oxides, used alone or as a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, alumina and silica-alumina. Very preferably, the oxide support essentially consists of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight, of alumina, relative to the total weight of said support. It preferably consists solely of alumina.

Preferably, the support of the first polishing hydrodesulfurization catalyst has a specific surface area of between 60 and 250 m²/g, preferably between 70 and 200 m²/g.

The total pore volume of the support of the first polishing hydrodesulfurization catalyst is generally between 0.3 cm³/g and 1.3 cm³/g, preferably of between 0.4 cm³/g and 1.1 cm³/g.

The support of the first polishing hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometry, platelets, of pellets, a compressed cylinder, crushed solids or any other shaping. Preferably, the support is in the form of beads with a diameter of 0.5 to 6 mm or in the form of cylindrical, trilobe or quadrilobe extrudates with a circumscribed diameter of 0.8 to 3 mm. More preferentially, the support is in the form of beads.

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

Step c) Second Polishing Hydrodesulfurization Step (FNS2)

The residual sulfur compounds on conclusion of step b) are essentially refractory sulfur compounds. Step c) of the process according to the invention essentially consists in converting at least a portion of the refractory sulfur compounds contained in the effluent resulting from step b), such as thiophene compounds, into saturated compounds, for example into thiophanes (or thiacyclopentanes) or into mercaptans, then in at least partially hydrogenolyzing these sulfur compounds to form H₂S.

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

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

The amount of catalyst employed in each reactor is generally such that the ratio of the volume flow rate of gasoline to be treated, expressed in m³ per hour under standard conditions (15° C., 0.1 MPa), per m³ of catalyst bed (also called hourly space velocity) is between 1 and 40 h⁻¹ and preferably between 2 and 20 h⁻¹.

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

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

The second polishing hydrodesulfurization catalyst of step c) comprises an active phase comprising, preferably consisting of, at least one group VIB metal and at least one group VIII metal at least partly in sulfide form, optionally phosphorus, and an oxide support, as described below.

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

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

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

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

Optionally, the catalyst can additionally have a phosphorus content generally of between 0.3% and 10% by weight of P₂O₅ relative to the total weight of catalyst, preferably between 0.3% and 5% by weight, very preferably between 0.5% and 3% by weight. Furthermore, when phosphorus is present, the phosphorus/(group VIB metal) molar ratio is generally between 0.1 and 0.7 mol/mol, preferably between 0.2 and 0.6 mol/mol.

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

The total pore volume of the catalyst of step c) is generally between 0.3 cm³/g and 1.3 cm³/g, preferably between 0.4 cm³/g and 1.1 cm³/g.

The oxide support of the hydrodesulfurization catalyst is typically a porous solid chosen from the group consisting of: alumina, silica, silica-alumina or else titanium or magnesium oxides, used alone or as a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, alumina and silica-alumina. Very preferably, the oxide support essentially consists of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight, of alumina, relative to the total weight of said support. It preferably consists solely of alumina.

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

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

The total pore volume of the support of the second polishing hydrodesulfurization catalyst is generally between 0.3 cm³/g and 1.3 cm³/g, preferably between 0.4 cm³/g and 1.1 cm³/g.

The support of the second polishing hydrodesulfurization catalyst can be in the form of beads, extrudates of any geometry, platelets, pellets, a compressed cylinder, crushed solids or any other shaping. Preferably, the support is in the form of beads with a diameter of 0.5 to 6 mm or in the form of cylindrical, trilobe or quadrilobe extrudates with a circumscribed diameter of 0.8 to 3 mm. More preferentially, the support is in the form of beads.

In a preferred variant, the catalyst employed during step c) is the same as that employed during step a).

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

Implementation of Steps b) and c) of the Process

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

When steps b) and c) of the process according to the invention are carried out in two different reactors, step b) can be carried out in a first polishing hydrodesulfurization reactor containing the second reaction section traversed by the partially desulfurized effluent resulting from step a), and then step c) can be carried out in the second polishing hydrodesulfurization reactor containing the third reaction section, placed downstream of said first reactor.

When steps b) and c) of the process according to the invention are carried out in a single reactor, step b) is carried out in a first region containing the second reaction section and step c) is carried out in a second region containing the third reaction section downstream of the first region.

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

Step d): Separation of the H₂S [Optional]

This separation step d) is carried out in order to separate the excess hydrogen and also the H₂S formed during steps a), b) and c). Any method known to a person skilled in the art can be envisaged.

According to a first embodiment, after steps a), b) and c), the third desulfurized effluent resulting from step c) is cooled to a temperature generally of less than 80° C. in order to condense the hydrocarbons. The gas and liquid phases are subsequently separated in a separation drum. The liquid fraction, which contains the desulfurized gasoline and also a fraction of dissolved H₂S, is sent to a stabilizer column or debutanizer. This column separates a top cut, consisting essentially of residual H₂S and of hydrocarbon compounds having a boiling point lower than or equal to that of butane, and a bottom cut stripped of H₂S, referred to as stabilized gasoline, containing the compounds having a boiling point greater than that of n-butane.

According to a second embodiment, after the condensation step, the liquid fraction which contains the desulfurized gasoline and also a a fraction of dissolved H₂S is sent to a stripping section, while the gaseous fraction, consisting mainly of hydrogen and H₂S, is sent to a purification section. The stripping can be carried out by heating the hydrocarbon fraction, alone or with an injection of hydrogen or steam, in a distillation column in order to extract, at the top, the light compounds which were entrained by dissolution in the liquid fraction and also the dissolved residual H₂S. The temperature of the stripped gasoline recovered at the column bottom is generally between 120° C. and 250° C.

Preferably, the separation step d) is carried out in a stabilizer column or debutanizer. This is because a stabilizer column makes it possible to separate the H₂S more efficiently than a stripping section.

Step e) is preferably carried out in order for the sulfur in the form of H₂S remaining in the desulfurized gasoline to represent less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.

Preparation of the Catalysts

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

Use may be made, by way of example, among the sources of molybdenum, of the oxides and hydroxides, molybdic acids and salts thereof, in particular the ammonium salts, such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H₃PMi₁₂O₄₀), and salts thereof, and optionally silicomolybdic acid (H₄SiMo₁₂O₄₀) and salts thereof. The sources of molybdenum can also be any heteropolycompound of Keggin, lacunary Keggin, substituted Keggin, Dawson, Anderson or Strandberg type, for example. Use is preferably made of molybdenum trioxide and the heteropolycompounds of Keggin, lacunary Keggin, substituted Keggin and Strandberg type.

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

The nickel precursors which can be used are advantageously chosen from the oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example. Use is preferably made of nickel hydroxide and nickel hydroxycarbonate.

The tungsten precursors which can be used are also well known to a person skilled in the art. For example, use may be made, among the sources of tungsten, of the oxides and hydroxides, tungstic acids and salts thereof, in particular the ammonium salts, such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and salts thereof, and optionally silicotungstic acid (H₄SiW₁₂O₄₀) and salts thereof. The sources of tungsten can also be any heteropolycompound of Keggin, lacunary Keggin, substituted Keggin or Dawson type, for example. Use is preferably made of the oxides and the ammonium salts, such as ammonium metatungstate, or the heteropolyanions of Keggin, lacunary Keggin or substituted Keggin type.

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

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

Following the maturation step, the catalyst precursor obtained can undergo a heat treatment. The aim of this treatment is generally to transform the molecular precursors of the elements into the oxide phase. It is in this case an oxidizing treatment, but a simple drying of the catalyst can also be carried out.

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

In the case of an oxidizing treatment, also referred to as calcination, said treatment is generally carried out under air or under 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 of between 0.5 hour and 24 hours, preferably for a period from 0.5 hour to 12 hours, and even more preferably for a period from 0.5 hour to 10 hours.

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 sulfurization. This activation phase is carried out by methods well known to a person skilled in the art, and advantageously under a sulfo-reductive atmosphere in the presence of hydrogen and of hydrogen sulfide. The hydrogen sulfide can be used directly or generated by a sulfide agent (such as dimethyl disulfide).

Description of the Sulfidation of the Catalysts

Before being brought into contact with the feedstock to be treated in a process for the hydrodesulfurization of gasolines, the catalysts used in steps, optionally a0), a), b) and c) of the process according to the invention generally undergo a sulfidation step. The sulfidation is preferably carried out in a sulforeducing medium, that is to say in the presence of H₂S and of hydrogen, in order to convert the metal oxides into sulfides, such as, for example, MoS₂, CogS₈ or NisS₂. The sulfidation is carried out by injecting, onto the catalyst, a stream containing H₂S and hydrogen, or else a sulfur compound capable of decomposing to H₂S in the presence of the catalyst and hydrogen. Polysulfides, such as dimethyl disulfide (DMDS), are H₂S precursors commonly used to sulfide the catalysts of steps, optionally a0), a), b) and c). The sulfur can also originate from the feedstock. The temperature is adjusted in order for the H₂S to react 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 temperatures of between 200° C. and 600° C. and more preferentially between 300° C. and 500° C.

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

${\left(S/\mathrm{metal}\right)}_{\mathrm{catayst}}\ge 0.6\times {\left(S/\mathrm{metal}\right)}_{\mathrm{theoretical}}$

• • in which:
  • (S/metal)_catalyst is the molar ratio of the sulfur(S) to the metal which are present on the catalyst
  • (S/metal)_theoretical is the molar ratio of the sulfur to the metal corresponding to the complete sulfidation of the metal to give sulfide.

This theoretical molar ratio varies according to the metal under consideration:

${\left(S/\mathrm{Co}\right)}_{\mathrm{theoretica1}}=1$ ${\left(S/\mathrm{Ni}\right)}_{\mathrm{theoretical}}=1$ ${\left(S/\mathrm{Mo}\right)}_{\mathrm{theoretical}}=2/1$ ${\left(S/W\right)}_{\mathrm{theoretical}}=2/1$

When the catalyst used in step, optionally a0), a) or in step c), comprises several metals, the molar ratio of the sulfur present on the catalyst to the combined metals also has to be at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of each metal to give sulfide, the calculation being carried out in proportion to the relative molar fractions of each metal.

The examples below illustrate the invention without limiting the scope thereof.

EXAMPLES

The analytical methods used to characterize the feedstocks and effluents are as follows:

• • sulfur content according to the ASTM D2622 method for contents above 10 ppm S and the ISO 20846 method for contents below 10 ppm S;
  • content of mercaptans according to the ASTM D3227 method;
  • content of olefins based on gas chromatography analysis according to the ASTM D6733 method.

Example 1: Preparation of Catalyst A

A support A′ is provided, which support is composed of alumina in the form of beads with a particle size of between 2 and 4 mm, and having a specific surface area of 139 m²/g and a pore volume of 0.97 ml/g.

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

The final element composition of catalyst A, expressed in the form of oxides and relative to the weight of dry catalyst, is then the following: MoO₃=10.0±0.2 wt % and CoO=3.0±0.1 wt %. The Co/Mo molar ratio is 0.60 mol/mol.

The specific surface of catalyst A is 124 m²/g.

Example 2: Preparation of Catalyst B

A support B′ is provided, which support is identical to the support A′.

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

The final element composition of catalyst B, expressed in the form of oxides and relative to the weight of dry catalyst, is then the following: NiO=17.9±0.3 wt %.

The specific surface of catalyst B is 114 m²/g.

Example 3: Use of the Catalysts in a Gasoline Desulfurization Process

Example 3 aims to show the advantage 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% by weight of olefins and 600 ppmS of total sulfur is subjected to a treatment in several steps:

• • a step of selective hydrodesulfurization (HDS) in an adiabatic reactor using catalyst A. The operating conditions of the step of one-step hydrodesulfurization of the gasoline feedstock are as follows: HSV=3 h⁻¹, P=2.0 MPa. A stream of pure hydrogen is added to feedstock at the reactor inlet such that H₂/HC=250 Nm³/m³. The effluent is sent directly to the reactor of the second step;
  • a first polishing hydrodesulfurization (FNS1) step in an adiabatic reactor using catalysts A or B. Only the effluent from the first step is treated in this second step. The pressure of the first polishing hydrodesulfurization step is set at 2.0 MPa. The reactor inlet temperature is always set at 35° C. higher than the temperature of the first effluent leaving the selective hydrodesulfurization step;
  • optionally, a second polishing hydrodesulfurization (FNS2) step in an adiabatic reactor using catalysts A or B. Only the effluent from the preceding step is treated in this second polishing hydrodesulfurization step. The pressure of the second polishing hydrodesulfurization step is set at 2.0 MPa. The reactor inlet temperature is equal to the temperature of the effluent leaving the preceding step.

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

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

TABLE 1
	Implementation	Implementation	Implementation
	not in	not in	not in	Implementation
	accordance	accordance	accordance	in accordance
	with the	with the	with the	with the
Implementation	invention	invention	invention	invention
HDS
Catalyst	A	A	A	A
HSV (h−1)	3	3	3	3
Reactor inlet temperature (° C.)	253	220	240	237
Reactor outlet temperature (° C.)	289	224	260	255
Mean temperature (° C.)	271	222	250	246
FNS1
Catalyst	B	A	A	B
HSV (h−1)	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	—	—	B	A
HSV (h−1)	—	—	4	12
Reactor inlet temperature (° C.)	—	—	317	291
Reactor outlet temperature (° C.)	—	—	319	310
Average temperature (° C.)			318	301
Overall performance
FNS1/FNS2 ratio (vol %)	—	—	25/75	75/25
Olefins at inlet (wt %)	25.0	25.0	25.0	25.0
Olefins at outlet (wt %)	16.5	15.6	15.3	16.7
Overall olefin hydrogenation (%)	34.0	37.6	38.8	33.2
Sulfur content at inlet (ppm wt S)	600	600	600	600
Sulfur content at outlet (ppm wt S)	10	10	10	10

The “average temperature” of the HDS, FNS1 or FNS2 step corresponds to the weight-average bed temperature (WABT), which is well known to a person skilled in the art. The average temperature is advantageously determined as a function of the catalytic systems, of the items of equipment, of the configuration thereof, that are used. The average temperature (or WABT) is calculated as follows:

$\mathrm{WABT}=\left(T_{\mathrm{entr}\stackrel{\prime }{e}e}+T_{\mathrm{sortie}}\right)/2$

with T_inlet: the temperature of the stream at the inlet of the reaction section and T_outlet: the temperature of the effluent at the outlet of the reaction section. Unless otherwise indicated, the “average temperature” of a reaction section is given under cycle start conditions.

Claims

1. A process for treating a gasoline containing sulfur compounds and olefins, the process comprising at least the following steps:

a) the gasoline, hydrogen and a hydrodesulfurization catalyst comprising an active phase comprising a group VIB metal and a group VIII metal at least partly in sulfide form, and an oxide support, are brought into contact in a first reaction section at a temperature of between 200° C. and 350° C., at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 h−1 and 20 h−1 and a ratio of the hydrogen flow rate, expressed in normal m3 per hour, to the volume flow rate of feedstock to be treated, expressed in m3 per hour under standard conditions, of between 10 Nm3/m3 and 1000 Nm3/m3, in order to obtain a first partially desulfurized effluent;

b) without separating the H2S formed in step a), the first partially desulfurized effluent obtained on conclusion of step a) and a first polishing hydrodesulfurization catalyst comprising an active phase consisting of a group VIII metal at least partly in sulfide form, and an oxide support, are brought directly into contact in a second reaction section at a temperature of between 250° C. and 400° C., at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 h−1 and 40 h−1, in order to obtain a second partially desulfurized effluent;

c) without separating the H2S formed in step b), the second partially desulfurized effluent obtained on conclusion of step b) and a second polishing hydrodesulfurization catalyst comprising an active phase comprising at least one group VIB metal and at least one group VIII metal at least partly in sulfide form, and an oxide support, are brought directly into contact in a third reaction section at a temperature of between 250° C. and 400° C., at a pressure of between 0.2 MPa and 5 MPa, with an hourly space velocity of between 1 h−1 and 40 h−1, in order to obtain a third partially desulfurized effluent.

2. The process as claimed in claim 1, wherein said second reaction section containing the first polishing hydrodesulfurization catalyst occupies a volume V1, and said third polishing hydrodesulfurization reaction section containing the second polishing hydrodesulfurization catalyst occupies a volume V2, the distribution of the volumes V1/V2 being between 90 vol %/10 vol % and 10 vol %/90 vol % respectively of said second and third reaction section.

3. The process as claimed in claim 1, wherein the catalyst of step a) and/or of step c) comprises a content of group VIII metal of between 0.1% and 10% by weight of oxide of the group VIII metal relative to the total weight of the catalyst, and a content of group VIB metal of between 1% and 20% by weight of oxide of the group VIB metal relative to the total weight of the catalyst.

4. The process as claimed in claim 1, wherein the catalyst of step a) and/or of step c) comprises alumina and an active phase comprising cobalt and molybdenum, said catalyst containing a content by weight, relative to the total weight of catalyst, of cobalt oxide, in CoO form, of between 0.1% and 10%, a content by weight, relative to the total weight of catalyst, of molybdenum oxide, in MoO3 form, of between 1% and 20%, and a cobalt/molybdenum molar ratio of between 0.1 and 0.8 mol/mol.

5. The process as claimed in claim 1, wherein the catalyst of step a) and/or of step c) further comprises phosphorus, said catalyst contains a content by weight, relative to the total weight of the catalyst, of phosphorus oxide, in P2O5 form, of between 0.3% and 10% by weight.

6. The process as claimed in claim 1, wherein the catalyst of step a) and/or of step c) has a specific surface area of between 60 and 250 m2/g.

7. The process as claimed in claim 1, wherein the catalysts of steps a) and c) are identical.

8. The process as claimed in claim 1, wherein the catalyst of step b) comprises a content of group VIII metal of between 5% and 65% by weight of oxide of the group VIII metal relative to the total weight of the catalyst.

9. The process as claimed in claim 1, wherein the catalyst of step b) comprises an alumina support and an active phase consisting of nickel, said catalyst containing a content by weight, relative to the total weight of catalyst, of nickel oxide, in NiO form, of between 5% and 65% by weight.

10. The process as claimed in claim 1, wherein the catalyst of step b) has a specific surface area of between 60 and 250 m2/g.

11. The process as claimed in claim 1, wherein, before step a), the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst in order to selectively hydrogenate the diolefins contained in said gasoline to give olefins.

12. The process as claimed in claim 1, wherein steps b) and c) are carried out in one and the same reactor.

13. The process as claimed in claim 1, wherein the temperature of steps b) and c) is higher than the temperature of step a).

14. The process as claimed in claim 1, wherein the gasoline is a catalytic cracking gasoline.