WO2023117531A1

WO2023117531A1 - Method for treating a petrol containing sulphur compounds

Info

Publication number: WO2023117531A1
Application number: PCT/EP2022/085356
Authority: WO
Inventors: Sophie COUDERC; Marie DEHLINGER; Adrien Gomez; Damien Hudebine; Marie-Claire Marion; Antoine Fecant
Original assignee: IFP Energies Nouvelles
Priority date: 2021-12-20
Filing date: 2022-12-12
Publication date: 2023-06-29
Also published as: AR128005A1; FR3130834A1; AU2022417857A1

Abstract

Disclosed is a method for treating a petrol containing sulphur compounds, olefins and diolefins, the method comprising at least the following steps: (a) bringing the petrol into contact, in at least one reactor, with hydrogen and a hydrodesulphurisation catalyst comprising an oxide support and an active phase comprising a group VI B metal and a group VIII metal; (b) bringing the effluent obtained on completion of step (a) into contact, in at least one reactor, with hydrogen and a hydrodesulphurisation catalyst comprising an oxide support and an active phase consisting of at least one group VIII metal; (c) transferring the effluent obtained on completion of step (b) to a separation vessel operating at a pressure of between 1.0 and 2.0 MPa to obtain a gaseous fraction containing H2S and hydrogen and a liquid fraction containing the desulphurised petrol and a dissolved H2S fraction; (d) transferring said liquid fraction obtained on completion of step (c) to a stabilising column so as to obtain, at the top, a stream comprising the residual H2S and C4-hydrocarbon compounds and, at the bottom, a stabilised petrol; (e) at least partially recycling the gaseous fraction obtained on completion of step (c) to at least one of steps (a) and/or (b).

Description

Process for treating gasoline containing sulfur compounds

Field of invention

The present invention relates to a method for producing gasoline with a low sulfur and mercaptan content.

State of the art

The production of gasolines meeting the new environmental standards requires a significant reduction in their sulfur content.

It is also known that conversion gasolines, and more particularly those originating from catalytic cracking, which can represent 30 to 50% of the gasoline pool, have high mono-olefin and sulfur contents.

The sulfur present in the gasolines is for this reason attributable, at nearly 90%, to the gasolines resulting from the processes of catalytic cracking, which will be called in the following gasolines of FCC (Fluid Catalytic Cracking according to the Anglo-Saxon terminology, that can be translated by catalytic cracking in a fluidized bed). FCC gasolines therefore constitute the preferred feedstock for the process of the present invention.

Among the possible ways to produce fuels with a low sulfur content, the one that has been very widely adopted consists in specifically treating sulfur-rich gasoline bases by catalytic hydrodesulphurization processes in the presence of hydrogen. Traditional processes desulfurize gasolines in a non-selective way by hydrogenating a large part of the mono-olefins, which generates a high loss in octane number and a high consumption of hydrogen. The most recent processes, such as the Prime G+ process (trademark), make it possible to desulphurize cracked gasolines rich in olefins, while limiting the hydrogenation of mono-olefins and consequently the loss of octane and the high consumption resulting hydrogen. Such processes are for example described in patent applications EP1077247 and EP1174485.

The residual sulfur compounds generally present in desulfurized gasoline can be separated into two distinct families: the unconverted refractory sulfur compounds present in the charge on the one hand, and the sulfur compounds formed in the reactor by so-called secondary recombination reactions. Among this last family of sulfur compounds, the majority compounds are the mercaptans resulting from the addition of the H2S formed in the reactor to the mono-olefins present in the charge.

Mercaptans, with the chemical formula R-SH, where R is an alkyl group, are also called recombinant mercaptans. Their formation or decomposition obeys the thermodynamic equilibrium of the reaction between mono-olefins and hydrogen sulfide to form recombinant mercaptans. An example is illustrated according to the following reaction:

The sulfur contained in the recombinant mercaptans generally represents between 20% and 80% by weight of the residual sulfur in the desulphurized gasolines.

The formation of recombination mercaptans is in particular described in patent US6231754 and patent application WO01/40409 which teach various combinations of operating conditions and of catalysts making it possible to limit the formation of recombination mercaptans.

Other solutions to the problem of the formation of recombination mercaptans are based on a treatment of partially desulfurized gasolines to extract said recombination mercaptans therefrom. Some of these solutions are described in patent applications WO02/28988 or WO01/79391.

Still other solutions are described in the literature for desulphurizing cracked gasolines using a combination of hydrodesulphurization and elimination of recombinant mercaptans by reaction to thioethers or disulphides (also called sweetening or “sweetening”) according to Anglo-Saxon terminology) (see for example US7799210, US6960291, US2007/114156, EP2861094).

Finally, patent application US2006/278567 discloses a process for the hydrodesulphurization of a cracked gasoline in two stages comprising intermediate separation stages making it possible in particular to limit the formation of mercaptans.

Objects of the invention

The subject of the present invention is a process for treating a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps: a) contacting in at least one reactor, the gasoline, hydrogen and a hydrodesulphurization catalyst comprising an oxide support and an active phase comprising a group VIB metal and a group VIII metal, at a temperature between 210 and 320°C, at a pressure between 1 .5 and 3 MPa, with a space velocity between 1 and 10 h ^-1 and a ratio between the hydrogen flow expressed in normal m ³ per hour and the feed rate to be treated expressed in m ³ per hour under standard conditions of between 100 and 600 Nm ³ /m ³ , so as to convert at least some of the sulfur compounds into H2S; b) in at least one reactor, the effluent from step a) without removal of the H ₂ S formed, is brought into contact with hydrogen and a hydrodesulphurization catalyst comprising an oxide support and a active phase consisting of at least one group VIII metal, at a temperature of between 280 and 400°C, at a pressure of between 1.0 and 3 MPa, with a space velocity of between 1 and 10 h ^-1 and a ratio between the hydrogen flow rate expressed in normal m ³ per hour and the feed flow rate to be treated expressed in m ³ per hour at standard conditions of between 100 and 600 Nm ³ /m ³ ; c) the effluent from step b) is sent to at least one separation drum operating at a pressure of between 1.0 and 2.0 MPa to obtain a gaseous fraction containing H2S and hydrogen and a liquid fraction containing the desulfurized gasoline and a fraction of dissolved H2S; d) said liquid fraction obtained in step c) is sent to a stabilization column to obtain at the top a stream comprising the residual H2S and C4- hydrocarbon compounds and at the bottom a stabilized gasoline; e) the gaseous fraction obtained at the end of step c) is recycled at least in part to at least one of steps a) and/or b).

The Applicant has identified, surprisingly, that in order to maintain a mercaptan and sulfur content in the hydrodesulfurized gasoline below 10 ppm by weight while limiting the loss of octane, the pressure in the hydrodesulfurization section (step a) of the process ) must be between 1.5 and 3 MPa and the pressure of the separation vessel must be within a specific range between 1.0 and 2.0 MPa. Indeed, if the pressure of the separation drum is less than 1.0 MPa, the purity of the hydrogen contained in the recycle gas fraction will decrease, thus causing a decrease in the ratio between the hydrogen flow rate and the flow rate of load to be processed. Thus, to maintain a content of mercaptans and sulfur of 10 ppm by weight, it is necessary to increase the temperature of the hydrodesulphurization stages, which has the effect of increasing the hydrogenation of the olefins of the gasoline cut, which will have has the effect of increasing octane loss. Furthermore, if the pressure of the separation drum is greater than 2.0 MPa, the purity of the hydrogen contained in the recycle gaseous fraction will certainly increase, but this leads to an increase in the pressure at the inlet of the reactors of hydrodesulfurization, and therefore, in addition to an overconsumption of hydrogen, will lead to an increase in the mercaptan content in the hydrodesulfurized gasoline and an increase in the loss of octane. According to one or more embodiments, the pressure of the separation drum of step c) is between 1.2 and 1.8 MPa.

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

According to one or more embodiments, the catalyst of stage a) comprises a molar ratio of metal from group VIII to metal from group VIB of the catalyst of between 0.1 and 0.8.

According to one or more embodiments, the catalyst of step a) comprises a specific surface of between 5 and 400 m ² /g.

According to one or more embodiments, the catalyst of step a) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight relative to the total weight of catalyst of cobalt oxide, in the CoO form, of between 0.1 and 10%, a content by weight relative to the total weight of catalyst of molybdenum oxide, in the MoOa form, of between 1 and 20%, a molar ratio cobalt/molybdenum of between 0.1 and 0.8, a content by weight relative to the total weight of catalyst of phosphorus oxide in the P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface between 50 and 250 m ² /g.

According to one or more embodiments, the catalyst of step b) comprises a group VIII metal content of between 1 and 60% by weight of group VIII metal oxide relative to the total weight of the catalyst.

According to one or more embodiments, the catalyst of step b) comprises a specific surface of between 5 and 400 m ² /g.

According to one or more embodiments, the catalyst of step b) consists of alumina and nickel, said catalyst containing a content by weight relative to the total weight of catalyst of nickel oxide, in NiO form, comprised between 5 and 20%, said catalyst having a specific surface between 30 and 180 m ² /g. According to one or more embodiments, the temperature of step b) is higher than the temperature of step a).

According to one or more embodiments, the temperature of step b) is higher by at least 5° C. than the temperature of step a).

According to one or more embodiments, before step a), a gasoline distillation step is carried out so as to split said gasoline into at least two light and heavy gasoline cuts and the heavy gasoline cut is treated in steps a ), b), c), d) and e).

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

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

List of Figures

Figure 1 is a schematic representation of an embodiment according to the invention. Figure 2 is a schematic representation of a method not in accordance with the invention.

detailed description

Definitions

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

By specific surface area is meant the BET specific surface area (SBET in m ² /g) determined by nitrogen adsorption in accordance with the ASTM D 3663-78 standard established from the BRUNAUER-EMMETT-TELLER method described in the periodical "The Journal of American Society', 1938, 60, 309. By total pore volume of the catalyst or of the support used for the preparation of the catalyst is meant the volume measured by intrusion with a mercury porosimeter according to standard ASTM D4284 at a maximum pressure of 4000 bar (400 MPa), using a surface tension of 484 dyne/cm and a contact angle of 140°, for example with an Autopore III model device from the Microméritics® brand.

The wetting angle was taken as equal to 140° by following the recommendations of the work “Engineering techniques, treatise on analysis and characterization”, pages 1050-1055, written by Jean Charpin and Bernard Rasneur. In order to obtain better precision, the value of the total pore volume corresponds to the value of the total pore volume measured by intrusion with a mercury porosimeter measured on the sample minus the value of the total pore volume measured by intrusion with a mercury porosimeter measured on the same sample for a pressure corresponding to 30 psi (about 0.2 MPa).

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

Description of the load

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

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

The sulfur content of gasoline cuts produced by catalytic cracking (FCC) depends on the sulfur content of the feed treated by the FCC, on the presence or not of a preprocessing of the FCC load, as well as the end point of the cut. Generally, the sulfur contents of an entire 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 higher than 200° C., the sulfur contents are often higher than 1000 ppm by weight, they can even in certain cases reach values of the order of 4000 to 5000 ppm by weight.

The feed treated by the process according to the invention can be a feed containing sulfur compounds in a content greater than 1000 ppm by weight of sulfur, and often greater than 1500 ppm.

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

Step a): Hydrodesulfurization

The hydrodesulfurization stage a) is implemented to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H ₂ S which is then eliminated in stage c).

The hydrodesulfurization step a) consists in bringing the gasoline to be treated into contact with hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization.

According to a preferred embodiment of the invention, step a) is implemented with the aim of carrying out hydrodesulphurization selectively, that is to say with a degree of hydrogenation of the mono-olefins of less than 80%, preferably less than 70% and very preferably less than 60%.

The temperature is generally between 210 and 320°C and preferably between 220 and 290°C. The temperature used must be sufficient to maintain the gasoline to be treated in the vapor phase in the reactor. In the case where stage a) of hydrodesulphurization is carried out in several reactors in series, the temperature of each reactor is generally higher by at least 5° C., preferably by at least 10° C. and very preferably of at least 30° C. to the temperature of the reactor which precedes it.

The operating pressure for this step is generally between 1.5 and 3 MPa.

The quantity of catalyst used in each reactor is generally such that the ratio between the flow rate of gasoline to be treated, expressed in m ³ per hour at standard conditions, per m ³ of catalyst (also called space velocity) is between 1 and 10:00 ^a.m. and preferably between 2 and 8:00 ^a.m.

The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m ³ per hour (Nm ³ /h) and the feed rate to be treated expressed in m ³ per hour at standard conditions (15° C., 0.1 MPa) is between 100 and 600 Nm ³ /m ³ , preferably between 200 and 500 Nm ³ /m ³ . Normal m ³ means the quantity of gas in a volume of 1 m ³ at 0°C and 0.1 MPa.

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

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

The catalyst used in step a) must have good selectivity with respect to hydrodesulphurization reactions compared to the hydrogenation reaction of olefins. The hydrodesulfurization catalyst of step a) comprises an oxide support and an active phase comprising a metal from group VIB and a metal from group VIII and optionally phosphorus and/or an organic compound as described below.

The group VIB metal present in the active phase of the catalyst is preferably chosen from molybdenum and tungsten. The group VIII metal present in the active phase of the catalyst is preferably chosen from cobalt, nickel and a 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 between 0.1 and 10% by weight of group VIII metal oxide relative to the total weight of the catalyst, preferably between 0.6 and 8% by weight, preferably between 0 6 and 7% by weight, very preferably between 1 and 6% by weight.

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

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

Optionally, the catalyst may also have a phosphorus content generally between 0.3 and 10% by weight of P ₂ Os relative to the total weight of catalyst, preferably between 0.3 and 5% by weight, very preferably between between 0.5 and 3% by weight. For example, the phosphorus present in the catalyst is combined with the metal from group VIB and optionally also with the metal from group VIII in the form of heteropolyanions.

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

Preferably, the catalyst is characterized by a specific surface of between 5 and 400 m ² /g, preferably of between 10 and 250 m ² /g, preferably of between 50 and 250 m ² /g. The specific surface is determined in the present invention by the BET method according to the ASTM D3663 standard, as described in the work Rouquerol F.; Rouquerol J.; Singh K. “Adsorption by Powders & Porous Solids: Principle, methodology and applications”, Academic Press, 1999, for example by means of an Autopore III™ model apparatus from the Micromeritics™ brand.

The total pore volume of the catalyst is generally between 0.4 cm ³ /g and 1.3 cm ³ /g, preferably between 0.6 cm ³ /g and 1.1 cm ³ /g. The total porous volume is measured by mercury porosimetry according to standard ASTM D4284 with a wetting angle of 140°, as described in the same work.

The tapped packing density (TRD) of the catalyst is generally between 0.4 and 0.8 g/mL, preferably between 0.4 and 0.7 g/mL. DRT measurement consists of introducing the catalyst into a test tube, the volume of which has been previously determined, then, by vibration, compacting it until a constant volume is obtained. The apparent density of the packed product is calculated by comparing the mass introduced and the volume occupied after packing.

The catalyst can be in the form of small-diameter, cylindrical or multilobed (trilobed, quadrilobed, etc.) extrudates, or spheres.

The catalyst oxide support is usually a porous solid chosen from the group consisting of: aluminas, silica, silica alumina or even titanium or magnesium oxides used alone or in a mixture with alumina or silica alumina . It is preferably chosen from the group consisting of silica, the family of transition aluminas and alumina silicas, very preferably, the oxide support consists essentially of alumina, that is to say that it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight of alumina. It preferably consists solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, that is to say which contains aluminas of theta, delta, kappa or alpha phase, alone or as a mixture and an amount of less than 20 % gamma, chi or eta phase alumina. The catalyst may also further comprise at least one organic compound containing oxygen and/or nitrogen and/or sulfur before sulfurization.

A very preferred embodiment of the invention corresponds to the implementation for step a) of a catalyst comprising alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight relative to the total weight of cobalt oxide catalyst, in CoO form, of between 0.1 and 10%, a content by weight relative to the total weight of molybdenum oxide catalyst, in MoOa form, between 1 and 20%, a cobalt/molybdenum molar ratio of between 0.1 and 0.8, a content by weight relative to the total weight of catalyst of phosphorus oxide in the P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface between 50 and 250 m ² /g. According to one embodiment, the active phase consists of cobalt and molybdenum.

According to another embodiment, the active phase consists of cobalt, molybdenum and phosphorus.

Step b): Finishing hydrodesulfurization

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

The so-called finishing hydrodesulphurization step is mainly implemented to reduce the content of recombinant mercaptans. Preferably, step b) is carried out at a higher temperature than that of step a). Indeed, by using a higher temperature in this step compared to the temperature of step a), the formation of olefins and of H ₂ S will be favored by the thermodynamic equilibrium. Stage b) also makes it possible to hydrodesulphurize the more refractory sulfur compounds.

The hydrodesulphurization step b) consists in bringing the effluent from step a) into contact, optionally with an addition of hydrogen, in one or more hydrodesulphurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulphurization .

The hydrodesulfurization step b) is carried out without significant hydrogenation of the olefins. The degree of hydrogenation of the olefins of the catalyst of the hydrodesulphurization stage b) is as a rule less than 5% and even more generally less than 2%.

The temperature of this step is generally between 280 and 400°C, more preferably between 290 and 380°C, and very preferably between 300 and 360°C. There temperature of this stage b) is generally higher by at least 5° C., preferably by at least 10° C. and very preferably by at least 30° C. than the temperature of stage a). The operating pressure for this step is generally between 1.0 and 3 MPa and preferably between 1.5 and 3 MPa.

Preferably, the flow rate of hydrogen is sustained and equal to the quantity injected in step a) minus the hydrogen consumed in step a). The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m ³ per hour (Nm ³ /h) and the feed flow rate to be treated expressed in m ³ per hour at standard conditions (15°C , 0.1 MPa) is between 100 and 600 Nm ³ /m ³ , preferably between 200 and 500 Nm ³ /m ³ .

The degree of desulfurization of step b), which depends on the sulfur content of the feed to be treated, is generally greater than 50% and preferably greater than 70% so that the product resulting from step b) contains less than 60 ppmw sulfur, preferably less than 40 ppmw sulfur, and most preferably less than 20 ppmw sulfur.

The hydrodesulfurization steps a) and b) can be carried out either in a single reactor containing the two catalysts, or in at least two different reactors. When steps a) and b) are carried out using two reactors, the latter two are placed in series, the second reactor fully treating the effluent at the outlet of the first reactor (without separation of the liquid and the gas between the first and the second reactor).

The catalyst of step b) is of a different nature and/or composition from that used in step a). The catalyst of stage b) is in particular a very selective hydrodesulphurization catalyst: it makes it possible to hydrodesulphurize without hydrogenating the olefins and therefore to maintain the octane number.

The catalyst which may be suitable for this step b) of the process according to the invention, without this list being limiting, is a catalyst comprising an oxide support and an active phase consisting of at least one group VIII metal, and preferably chosen from the group formed by nickel, cobalt, iron. These metals can be used alone or in combination. Preferably, the active phase consists of a group VIII metal, preferably nickel. Particularly preferably, the active phase consists of nickel. The group VIII metal content is between 1 and 60% by weight of group VIII metal oxide relative to the total weight of the catalyst, preferably between 5 and 30% by weight, very preferably between 5 and 20 % weight.

Preferably, the catalyst is characterized by a specific surface of between 5 and 400 m ² /g, preferably between 10 and 250 m ² /g, preferably between 20 and 200 m ² /g, very preferably between 30 and 180 m ² /g. The specific surface is determined in the present invention by the BET method according to the ASTM D3663 standard, as described in the work Rouquerol F.; Rouquerol J.; Singh K. “Adsorption by Powders & Porous Solids: Principle, methodology and applications”, Academic Press, 1999, for example by means of an Autopore III™ model apparatus from the Micromeritics™ brand.

The pore volume of the catalyst is generally between 0.4 cm ³ /g and 1.3 cm ³ /g, preferably between 0.6 cm ³ /g and 1.1 cm ³ /g. The total porous volume is measured by mercury porosimetry according to standard ASTM D4284 with a wetting angle of 140°, as described in the same work.

The tapped packing density (TRD) of the catalyst is generally between 0.4 and 0.8 g/mL, preferably between 0.4 and 0.7 g/mL.

DRT measurement consists of introducing the catalyst into a test tube, the volume of which has been previously determined, then, by vibration, compacting it until a constant volume is obtained. The apparent density of the packed product is calculated by comparing the mass introduced and the volume occupied after packing.

The catalyst oxide support is usually a porous solid chosen from the group consisting of: aluminas, silica, silica alumina or even titanium or magnesium oxides used alone or in a mixture with alumina or silica alumina . It is preferably chosen from the group consisting of silica, the family of transition aluminas and alumina silicas, very preferably, the oxide support consists essentially of alumina, that is to say that it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, or even at least 90% by weight of alumina. It preferably consists solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, that is to say which contains aluminas of theta, delta, kappa or alpha phase, alone or as a mixture and an amount of less than 20 % gamma, chi or eta phase alumina.

A very preferred embodiment of the invention corresponds to the implementation for step b) of a catalyst consisting of alumina and nickel, said catalyst containing a content by weight relative to the total weight of nickel oxide catalyst, in NiO form, of between 5 and 20%, said catalyst having a specific surface area of between 30 and 180 m ² /g.

The catalyst of the hydrodesulfurization stage b) is characterized by a catalytic hydrodesulfurization activity generally comprised between 1% and 90%, preferentially comprised between 1% and 70%, and very preferably comprised between 1% and 50% the catalytic activity of the catalyst of the hydrodesulphurization step a).

Step c): Separation section

This stage is implemented in order to separate the excess hydrogen as well as the H2S formed during stages a) and b) and to extract the desulfurized gasoline meeting the sulfur and recombination mercaptan specifications while reducing losses of C5+ hydrocarbon compounds (octane).

According to step c) of the process according to the invention, the effluent from step b) is fractionated so as to produce a gaseous phase comprising hydrogen from steps a) and b), H ₂ S formed during steps a) and b) and light compounds C1 to C4, and desulfurized gasoline.

More particularly, after the hydrodesulfurization step b), the gasoline is advantageously cooled to a temperature generally below 80° C., preferably below 60° C. in order to condense the hydrocarbons. The gas and liquid phases are then separated in a separation drum. The temperature of the separation vessel is generally between 20 and 80°C, preferably between 25 and 65°C. The pressure of the separator drum is set between 1.0 and 2.0 MPa, preferably between 1.2 and 1.8 MPa.

Stage c) is preferably implemented so that the sulfur in the form of H ₂ S remaining in the effluent from stage b) represents less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.

Step d): Stabilization

The liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved H ₂ S is sent to a stabilization column or debutaniser. This column separates a top cut essentially consisting of residual H ₂ S and hydrocarbon compounds having a boiling point less than or equal to that of butane (C4-) and a bottom cut freed from H ₂ S which corresponds to desulfurized and stabilized gasoline containing compounds having a boiling point higher than that of butane. The fact of using a separator drum operating advantageously at low temperature to eliminate the hL, the H2S and the light gases, followed by a stabilization column to eliminate the dissolved H2S and produce the stabilized desulfurized gasoline makes it possible to reduce the temperature at the bottom of the stabilization column compared to a separation of I' H2 and H2S directly in a stabilization column.

The stabilization column generally operates at a pressure between 0.1 and 2 MPa and preferably between 0.2 and 1 MPa.

The number of theoretical plates of this separation column is generally between 10 and 50 and preferably between 20 and 40.

The reflux rate, expressed as being the ratio of the liquid traffic in the column divided by the flow rate of distillate expressed in kg/h, is generally less than unity and preferably less than 0.5.

The desulfurized and stabilized heavy gasoline produced by the process according to the invention is advantageously used as a base for the formulation of a gasoline fuel.

Step e): Recycling

According to the invention, at least part of the gaseous fraction separated by the separating drum in step c) is recycled to steps a) and/or b), preferably to step b). Advantageously, said gaseous fraction is sent beforehand to an amine absorber or a washing column operating at low pressure (typically at 1.5 MPa), in order to eliminate at least part of the H ₂ S.

Description of catalyst preparation and sulfurization

The preparation of the catalysts of stages a) and b) is known and generally comprises a stage of impregnation of metals from group VIII and from group VIB when it is present, and optionally phosphorus and/or the organic compound on the support of oxide, followed by drying, then optional calcination to obtain the active phase in their oxide forms. Before its use in a hydrodesulfurization process of an olefinic gasoline cut containing sulfur, the catalysts are generally subjected to sulfurization in order to form the active species as described below.

The impregnation step can be carried out either by slurry impregnation, or by excess impregnation, or by dry impregnation, or by any other means known to those skilled in the art. The impregnation solution is chosen so as to be able to dissolve the metal precursors in the desired concentrations. By way of example, among the sources of molybdenum, use may be made of oxides and hydroxides, molybdic acids and their salts, in particular ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PM012O40), and their salts, and optionally silicomolybdic acid (H ₄ SiMoi ₂ 04o) and its salts. The sources of molybdenum can also be any heteropolycompound of Keggin, lacunary Keggin, substituted Keggin, Dawson, Anderson, Strandberg type, for example. Preferably, molybdenum trioxide and the heteropolycompounds of Keggin, lacunary Keggin, substituted Keggin and Strandberg type are used.

The tungsten precursors which can be used are also well known to those skilled in the art. For example, among the sources of tungsten, it is possible to use oxides and hydroxides, tungstic acids and their salts, in particular ammonium salts such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and their salts, and optionally silicotungstic acid (H ₄ SiWi ₂ 04o) and its salts. The tungsten sources can also be any heteropolycompound of Keggin, lacunary Keggin, substituted Keggin, Dawson type, for example. Preferably, ammonium oxides and salts are used, such as ammonium metatungstate or heteropolyanions of Keggin, lacunary Keggin or substituted Keggin type.

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

The nickel precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example. The preferred phosphorus precursor is orthophosphoric acid H3PO4, but its salts and esters such as ammonium phosphates are also suitable. The phosphorus can also be introduced at the same time as the element(s) of group VI B in the form of heteropolyanions of Keggin, lacunary Keggin, substituted Keggin or of the Strandberg type.

After the impregnation step, the catalyst is generally subjected to a drying step at a temperature below 200° C., advantageously between 50° C. and 180° C., preferably between 70° C. and 150° C., very preferably between 75°C and 130°C. The drying step is preferably carried out under an inert atmosphere or under an atmosphere containing oxygen. The drying step can be carried out by any technique known to those skilled in the art. It is advantageously carried out at atmospheric pressure or at reduced pressure. Preferably, this step is carried out at atmospheric pressure. It is advantageously carried out in a traversed bed using air or any other hot gas. Preferably, when the drying is carried out in a fixed bed, the gas used is either air or an inert gas such as argon or nitrogen. Very preferably, the drying is carried out in a traversed bed in the presence of nitrogen and/or air. Of preferably, the drying step lasts between 5 minutes and 15 hours, preferably between 30 minutes and 12 hours.

According to a variant of the invention, the catalyst has not undergone calcination during its preparation, that is to say that the impregnated catalytic precursor has not been subjected to a heat treatment step at a temperature higher than at 200° C. under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not. According to another variant of the invention, preferred, the catalyst has undergone a calcination step during its preparation, that is to say that the impregnated catalytic precursor has been subjected to a heat treatment step at a temperature between 250 and 1000° C. and preferably between 200 and 750° C., for a duration typically comprised between 15 minutes and 10 hours, under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.

Before being brought into contact with the feedstock to be treated in a gasoline hydrodesulfurization process, the catalysts of the process according to the invention generally undergo a sulfurization step. The sulfurization is preferably carried out in a sulphur-reducing medium, that is to say in the presence of H ₂ S and hydrogen, in order to transform the metal oxides into sulphides such as, for example, M0S2, CogSs or Ni ₃ S2. Sulfurization is carried out by injecting onto the catalyst a stream containing H2S and hydrogen, or else a sulfur compound capable of decomposing into H2S in the presence of the catalyst and hydrogen. Polysulphides such as dimethyldisulphide (DMDS) are H ₂ S precursors commonly used to sulphide catalysts. The sulfur can also come from the filler. The temperature is adjusted so that the H ₂ S reacts with the metal oxides to form metal sulphides. This sulfurization can be carried out in situ or ex situ (inside or outside the reactor) of the reactor of the process according to the invention at temperatures between 200 and 600°C, and more preferably between 300 and 500°C.

The sulfurization rate of the metals constituting the catalysts is at least equal to 60%, preferably at least equal to 80%. The sulfur content in the sulfurized catalyst is measured by elemental analysis according to ASTM D5373. A metal is considered to be sulfurized when the overall sulfurization rate defined by the molar ratio between the sulfur (S) present on the catalyst and said metal is at least equal to 60% of the theoretical molar ratio corresponding to the total sulfurization of the metal(s) considered. The overall sulfurization rate is defined by the following equation: (S/metal) _C atalyser — 0.6 X (S/metal)theoretical in which:

(S/metal) _catalyst is the molar ratio between sulfur (S) and metal present on the catalyst (S/metal) theoretical _t is the molar ratio between sulfur and metal corresponding to the total sulphidation of the metal to sulphide.

This theoretical molar ratio varies according to the metal considered:

(S/Fe)theoretical ^— 1

(S/Co)theoretical ^— 1

(S/Ni)theoretical ^— 1

(S/Mo)theoretical ^— 2/1

(S/W)theoretical =2/1

When the catalyst comprises several metals, the molar ratio between the S present on the catalyst and all the metals must also be at least equal to 60% of the theoretical molar ratio corresponding to the total sulphidation of each metal to sulphide, the calculation being carried out in proportion to the relative molar fractions of each metal.

Schemes that can be implemented in the context of the invention

Different schemes can be implemented in order to produce, at lower cost, a desulfurized gasoline with a reduced mercaptan content. The choice of the optimal scheme depends in fact on the characteristics of the species to be treated and produced as well as the constraints specific to each refinery.

The diagrams described below are given by way of non-limiting illustration.

According to a first variant, a step of distillation of the gasoline to be treated is carried out in order to separate two cuts (or fractions), namely a light cut and a heavy cut and the heavy cut is treated according to the method of the invention. . The light cut generally has a boiling temperature range below 100°C and the heavy cut a temperature range above 65°C. This first variant has the advantage of not hydrotreating the light cut which is rich in olefins and generally low in sulphur, which makes it possible to limit the loss of octane by hydrogenation of the olefins.

According to a second variant, the gasoline to be treated is subjected before the process according to the invention to a preliminary step consisting of a selective hydrogenation of the diolefins present in the feed, as described in patent application EP 1077247.

The gasoline to be treated is treated beforehand in the presence of hydrogen and a selective hydrogenation catalyst so as to at least partially hydrogenate the diolefins and carry out a reaction of weighting down a part of the light mercaptan compounds (RSH) present in the thioether charge, 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 diolefins and for the weighting of light mercaptans. The reaction of selective hydrogenation of the diolefins and weighting of the light mercaptans is preferably carried out on a sulfur catalyst comprising at least one element from group VIII and optionally at least one element from group VIB and an oxide support. The group VIII element is preferably chosen from nickel and cobalt and in particular nickel. The element of group VIB, when it is present, is preferably chosen from molybdenum and tungsten and very preferably molybdenum.

The catalyst oxide support is preferably chosen from alumina, nickel aluminate, silica, silicon carbide, or a mixture of these oxides. Preferably, alumina is used and even more preferably, high purity alumina. According to a preferred embodiment, the selective hydrogenation catalyst contains nickel at a content by weight of nickel oxide, in the form of NiO, of between 1 and 12%, and molybdenum at a content by weight of molybdenum oxide , in MoOa 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 sulfurization rate of the metals constituting the catalyst is preferably greater than 60%.

During the optional selective hydrogenation step, the gasoline is brought into contact with the catalyst at a temperature of between 50 and 250° C., and preferably between 80 and 220° C., and even more preferably between 90 and 200°C, with a liquid space velocity between 0.5 h ^-1 and 20 h ^-1 , the unit of the liquid space velocity being the liter of charge per liter of catalyst and per hour (L/L/h ). The pressure is between 0.4 and 5 MPa, preferably between 0.6 and 4 MPa and even more preferably between 1 and 3 MPa. The optional selective hydrogenation step is typically carried out with a ratio between the hydrogen flow rate expressed in normal m ³ per hour and the feed flow rate to be treated expressed in m ³ per hour at standard conditions of between 2 and 100 Nm ³ /m ³ , preferably between 3 and 30 Nm ³ /m ³ .

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

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

It should be noted that it is possible to carry out the stages 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 .

Other characteristics and advantages of the invention will now appear on reading the description which will follow, given solely by way of illustration and not limitation, and with reference to the appended FIG.

With reference to FIG. 1 and according to one embodiment of the process according to the invention, the gasoline to be treated is sent via line 1 and hydrogen via line 3 to a selective hydrogenation unit 2 (optional step ) to selectively hydrogenate diolefins and weigh down light mercaptans. The effluent with a low diolefin and mercaptan content is withdrawn from reactor 2 via line 4 and is sent to a fractionation column 5 (or splitter according to English terminology) configured to separate the gasoline into two cuts: a light gasoline cut 6 (or light gasoline) and a (first) heavy gasoline cut 7 which consists of the heavy fraction complementary to the light gasoline. The light cut point cut is generally performed at a temperature below 100°C, and the heavy cut point cut is generally performed at a temperature above 65°C. The final boiling point of the light cut is chosen so as to provide a light gasoline cut with a low sulfur content (total sulfur content typically less than 30 ppm by weight and preferably less than 10 ppm by weight) without requiring a step of subsequent hydrodesulfurization.

The heavy gasoline fraction is then sent via line 7 and hydrogen via line 8 and recycle gas via line 14a into the hydrodesulphurization unit 9 of step a). The hydrodesulfurization unit 9 of step a) is for example a reactor containing a supported hydrodesulfurization catalyst based on a metal from groups VIII and VI B in a fixed bed or in a fluidized bed, preferably a fixed bed reactor. The reactor is operated under operating conditions and in the presence of a hydrodesulphurization catalyst, as described above, to decompose the sulfur compounds and form hydrogen sulphide (H2S). During the hydrodesulfurization in step a), recombination mercaptans are formed by addition of H ₂ S formed on the olefins. The effluent from hydrodesulfurization unit 9 is then introduced into so-called finishing hydrodesulfurization unit 11 via line 10 without removing the H ₂ S formed. The recycle gas obtained is optionally sent after step of separation by line 14b in the finishing hydrodesulfurization unit 11. The hydrodesulfurization unit 11 of step b) is for example a reactor containing a hydrodesulfurization catalyst in a fixed bed or in a fluidized bed, preferably a fixed bed reactor is used. Unit 11 is operated at a higher temperature than unit 9 and in the presence of a selective catalyst comprising an oxide support and an active phase consisting of at least one group VIII metal to at least partially decompose recombination mercaptans into olefins and H ₂ S. It also makes it possible to hydrodesulfurize the more refractory sulfur compounds.

According to one or more embodiments, recycle gas is sent via line 14b into the so-called finishing hydrodesulfurization unit 11, and recycle gas is optionally sent via line 14a into the hydrodesulfurization unit 9 .

An effluent (gasoline) containing H ₂ S is withdrawn from said hydrodesulphurization reactor 11 via line 12. The effluent then undergoes a stage of elimination of rH ₂ S (stage c) which consists, in the mode of embodiment of Figure 1, to treat the effluent by condensation by introducing the effluent from step b) via line 12 into a separation drum 13 whose pressure is set so as to guarantee a content of mercaptans and sulfur in effluent less than 10 ppmw while maintaining optimum cycle time on the HDS catalyst. This separation drum 13 makes it possible to obtain a gaseous fraction containing H ₂ S and hydrogen via line 14 and to withdraw a liquid fraction. The gaseous fraction, after removal of the H ₂ S, is recycled to the hydrodesulfurization units 9 and/or 11 via lines 14a and/or 14b. The liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved H ₂ S is sent via line 15 to a stabilization column or debutanizer 16 in order to separate at the top of the column via line 17 a stream containing C4- hydrocarbons and residual H ₂ S and at the bottom of the column via line 18 a so-called stabilized gasoline containing compounds having a boiling point higher than that of butane and which has mercaptan and sulfur contents total less than 10 ppm by weight.

With the sequences proposed for the process according to the invention, it is possible to achieve high hydrodesulphurization rates while limiting the loss of olefins and consequently the decrease in the octane number. The process according to the invention also makes it possible to reduce the losses of C5+ hydrocarbon compounds.

The examples below illustrate the invention without limiting its scope. Examples

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

- density according to the NF EN ISO 12185 method;

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

- distillation according to the CSD method simulated distillation “CSD according to the ASTM2887 method;

- diolefin content, determined using the maleic anhydride index (MAV or "Maleic Anhydride Value" according to English terminology), according to the UOP 326 method.

Example 1 (non-compliant)

This example refers to Figure 2.

Table 1 gives the characteristics of an FCC gasoline treated by the process according to Figure 2.

The FCC gasoline (line 1) is treated in the selective hydrogenation reactor 2 in the presence of a catalyst A. The catalyst A is a catalyst of the NiMo on alumina type. The metal contents are respectively 7% by weight NiO and 11% by weight MoOa relative to the total weight of the catalyst, ie a Ni/Mo molar ratio of 1.2. The specific surface of the catalyst is 230 m ² /g. Prior to its use, catalyst A is sulfurized at atmospheric pressure in a sulfurization bench under an H2S/H2 mixture consisting of 15% by volume of FhS at 1 L/gh of catalyst and at 400° C. for two hours. This protocol makes it possible to obtain a sulfurization rate of over 80%.

The gasoline (line 1) is brought into contact with hydrogen (line 3) in a reactor which contains catalyst A. This stage of the process carries out the selective hydrogenation of the diolefins and the conversion (weighting down) of a part light mercaptan compounds (RSH) present in the filler. The diolefin content is directly proportional to the value of the MAV (maleic anhydride index or Maleic Anhydrid Value according to the Anglo-Saxon terminology). Diolefins are undesirable compounds because they are gum precursors in gasoline.

The operating conditions used in the selective hydrogenation reactor are: Temperature: 140°C, Total pressure: 2.5 MPa, H2 volume ratio added/gasoline charge: 5 normal liters of hydrogen per liter of gasoline at the conditions standard (flight/flight), liquid space velocity: 3 hr ¹ .

Table 1: Characteristics of feed 1 and selective hydrogenation effluent 4.

The effluent from the selective hydrogenation stage (line 4) with a low content of conjugated diolefins (MAV=0.6 mg/g) and a low content of light sulfur compounds (weighted down in the selective hydrogenation stage) is sent to a fractionation column 5 in order to separate a light gasoline at the top (line 6) and at the bottom of the column a first heavy gasoline cut (line 7). The characteristics of the light gasoline and of the first heavy gasoline cut are indicated in Table 2. As indicated in Table 2, the light gasoline obtained (line 6) has a low sulfur content (10 ppm by weight). The first heavy gasoline cut, which corresponds to around 72% by weight of the gasoline, has a high sulfur content (600 ppm) and requires additional treatment before being incorporated into the gasoline pool.

Table 2: Characteristics of cuts: Light gasoline and first cut heavy gasoline

The first heavy gasoline cut (line 7) is mixed with hydrogen (line 8) and treated in a selective hydrodesulphurization unit (9) which corresponds to a first hydrodesulphurization stage. The first hydrodesulphurization step is carried out in the presence of a CoMo-type catalyst supported on alumina, the metal contents being respectively 3% by weight of CoO and 10% by weight of MoOa, the specific surface of the catalyst is 135 m ² /g. Prior to its use, the catalyst is sulfurized at atmospheric pressure in a sulfurization bench under an H ₂ S/H ₂ mixture consisting of 15% by volume of H ₂ S at 1 L/gh of catalyst and at 400° C. for two hours. This protocol makes it possible to obtain a sulfurization rate of over 80%. The temperature is 270° C., the pressure is 2.1 MPa, the space velocity of the liquid (expressed in volume of liquid per volume of catalyst and per hour) is 3 h' ¹ , the ratio between the flow of hydrogen and the feed rate is 250 normal m ³ per m ³ under standard conditions. The reactor effluent (line 10) is then reheated in a furnace (not shown in the figure) then introduced into a second reactor (11) containing a so-called finishing catalyst. This finishing step is carried out in the presence of an Ni catalyst supported on alumina. The temperature is 324° C., the pressure is 1.8 MPa, the space velocity of the liquid (expressed in volume of liquid per volume of catalyst and per hour) is 3 h′ ¹ .

The effluent from reactor 11 (line 12) is sent to a stabilization or debutanizer column (16) operating at a pressure of 1.6 MPa, in order to separate at the top of the column via line 17 a stream containing H ₂ S, hydrogen, C4- hydrocarbons and at the bottom of the column, via line 18, a so-called stabilized heavy gasoline, the characteristics of which are illustrated in table 3. The loss in gasoline yield is presented in table 4.

Table 3: Characteristics of heavy gasoline after the first and second hydrodesulphurization stages

Table 4: Loss of C5+ compounds between the first heavy gasoline cut (line 7) and the gasoline obtained after the second hydrodesulphurization stage and after the stabilization stage (line 18)

The process as described in example 1 makes it possible to obtain a heavy gasoline with a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 35.7% by weight (relative).

Example 2 (compliant)

This example refers to the present invention according to Figure 1.

Unless otherwise indicated below, the same feed as in Example 1 is sent under the same operating conditions as Example 1 to the selective hydrogenation, hydrodesulphurization and finishing hydrodesulphurization sections.

The effluent from reactor 11 (line 12) is sent to a separation drum (13) after condensation at a temperature of 65° C. and a pressure of 1.6 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved hLS (line 15) and a gaseous fraction (line 14) essentially containing the hydrogen and the H ₂ S formed during steps a) and b), optionally with light C1 to C4 hydrocarbons. The liquid fraction (line 15) is then sent to a stabilization column (16) operating at a pressure of 0.7 MPa, in order to separate at the head of the column a cut of light compounds from the liquid fraction as well as the H2S (line 17) and a bottom section stripped of FLS, called stabilized gasoline (line 18).

The gas stream (line 14) is recycled to the first hydrodesulfurization stage (stage a), via line 14a.

The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in Table 5 below.

Table 5: Characteristics of the gaseous fraction (line 14) after gas/liquid separation of the reaction effluent 12

The characteristics of the gasoline cut obtained after stabilization (line 18) of the present invention are illustrated in Tables 6 and 7.

Table 6: Characteristics of heavy gasoline after the first and second hydrodesulphurization stages

Table 7: Loss of C5+ compounds between the first heavy gasoline cut (line 7) and the gasoline obtained after the second hydrodesulphurization stage and after the stabilization stage (line 18)

The process according to Example 2 makes it possible to obtain a heavy gasoline with a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulphurization step is 34.7% by weight (relative).

Example 3 (non-compliant)

The volume flow rate of the recycle gas loop is kept constant with respect to example 2, the ratio between the hydrogen flow rate and the charge rate is therefore a consequence and is 125 normal m ³ per m ³ of charge under standard conditions, the average temperature of the hydrodesulfurization reactor defined as the average between the inlet and outlet temperature of the reactor is 283°C, the pressure is 2.2 MPa, the space velocity of the liquid ( expressed in volume of liquid per volume of catalyst and per hour) is 3 h ^-1 .

The reactor effluent (line 10) is reheated in a furnace (not shown in the figure) then introduced into a second reactor (11) containing a so-called finishing catalyst. This finishing step is carried out at an average temperature of 338° C., a pressure of 1.9 MPa, and a space velocity of the liquid (expressed in volume of liquid per volume of catalyst and per hour) of 3 h′ ¹ .

The effluent from reactor 11 (line 12) is sent to a separation drum (13) after condensation at a temperature of 65° C. and a pressure of 0.8 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved hLS (line 15) and a gaseous fraction (line 14) essentially containing the hydrogen and the hLS formed during step b), optionally with light hydrocarbons C1 to C4 and C5+ hydrocarbons. The gaseous fraction is sent to an hLS separation section in an amine washing column (19) in order to obtain a recycle gas containing 10 ppm mol of H2S. The recycle gas is returned via line 14a, mixed with fresh hydrogen (line 8) to the selective hydrodesulphurization unit (9) after compression.

The liquid fraction (line 15) from the separator operating at 0.8 MPa is then sent to a stabilization column (16) operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds of the liquid fraction as well as the H ₂ S (line 17) and a bottom section freed from the H ₂ S, called stabilized gasoline (line 18).

The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in Table 8 below.

Table 8: Characteristics of the gaseous fraction (line 14) after gas/liquid separation of the reaction effluent 12

The characteristics of the gasoline cut obtained after stabilization (line 18) are illustrated in Table 9.

Table 9: Characteristics of heavy gasoline after the first and second hydrodesulphurization stages The loss of C5+ compounds in the gaseous fraction obtained after the separation stage (line 14) and in the gaseous cut after stabilization of the gasoline (line 17) is given in Table 10.

Table 10: Loss of C5+ compounds between the gasoline obtained after the second hydrodesulphurization stage (line 12) and after the stabilization stage (line 18)

The process according to Example 3 makes it possible to obtain a heavy gasoline with a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulphurization step is 37.4% by weight (relative). This octane loss is greater than in the example according to the invention. It also increases hydrogen consumption in hydrodesulfurization reactors.

Example 4 (non-compliant)

The volume flow rate of the recycle gas loop is kept constant compared to example 2, the ratio between the hydrogen flow rate and the charge rate is therefore a consequence and is 330 normal m ³ per m ³ of charge under standard conditions, the average temperature of the hydrodesulfurization reactor defined as the average between the inlet and outlet temperature of the reactor is 268°C, the pressure is 2.7 MPa, the space velocity of the liquid ( expressed in volume of liquid per volume of catalyst and per hour) is 3 h ^-1 .

The reactor effluent (line 10) is reheated in a furnace (not shown in the figure) then introduced into a second reactor (11) containing a so-called finishing catalyst. This finishing step is carried out at an average temperature of 320° C., a pressure of 2.4 MPa, and a space velocity of the liquid (expressed in volume of liquid per volume of catalyst and per hour) of 3 h′ ¹ .

The effluent from reactor 11 (line 12) is sent to a separation drum (13) after condensation at a temperature of 65° C. and a pressure of 2.1 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved hLS (line 15) and a gaseous fraction (line 14) essentially containing the hydrogen and the hLS formed during step b), optionally with light hydrocarbons C1 to C4 and C5+ hydrocarbons. The gaseous fraction is sent to an H2S separation section in an amine washing column (19) in order to obtain a recycle gas containing 10 ppm mol of H2S. The recycle gas is returned via line 14a, mixed with fresh hydrogen (line 8) to the selective hydrodesulphurization unit (9) after compression. The liquid fraction (line 15) from the separator operating at 2.1 MPa is then sent to a stabilization column (16) operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds of the liquid fraction as well as the H ₂ S (line 17) and a bottom section freed from the H ₂ S, called stabilized gasoline (line 18).

The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in Table 11.

Table 11: Characteristics of the gaseous fraction (line 14) after gas/liquid separation of the reaction effluent 12

The characteristics of the gasoline cut obtained after stabilization (line 18) are illustrated in Table 12.

Table 12: Characteristics of heavy gasoline after the first and second hydrodesulphurization stages The loss of C5+ compounds in the gaseous fraction obtained after the separation stage (line 14) and in the gaseous cut after stabilization of the gasoline (line 17) is given in Table 13.

Table 13: Loss of C5+ compounds between the gasoline obtained after the second hydrodesulphurization stage (line 12) and the gasoline obtained after the stabilization stage (line 18)

The process according to Example 4 makes it possible to obtain a heavy gasoline with a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 37.8% by weight (relative). This octane loss is greater than in the example according to the invention. It also increases hydrogen consumption in hydrodesulfurization reactors.

Example 5 (non-compliant)

The volume flow rate of the recycle gas loop is kept constant with respect to example 1, the ratio between the hydrogen flow rate and the charge rate is therefore 250 normal m ³ per m ³ of charge under standard conditions. , the average temperature of the hydrodesulphurization reactor is 272°C, the pressure is 3.6 MPa, the space velocity of the liquid (expressed in volume of liquid per volume of catalyst and per hour) is 3 IT 1

The reactor effluent (line 10) is then reheated in a furnace (not shown in the figure) then introduced into a second reactor (11) containing a so-called finishing catalyst. This finishing step is carried out at an average temperature of 330° C., a pressure of 3.3 MPa, and a space velocity of the liquid (expressed in volume of liquid per volume of catalyst and per hour) of 3 h′ ¹ .

The effluent from reactor 11 (line 12) is sent to a separation drum (13) after condensation at a temperature of 65° C. and a pressure of 1.6 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved hLS (line 15) and a gaseous fraction (line 14) essentially containing the hydrogen and the hLS formed during step b), optionally with light hydrocarbons C1 to C4 and C5+ hydrocarbons. The gaseous fraction is sent to an H2S separation section in an amine washing column (19) in order to obtain a recycle gas containing 10 ppm mol of H2S. The recycle gas is returned via line 14a, mixed with fresh hydrogen (line 8) to the selective hydrodesulphurization unit (9) after compression. The liquid fraction (line 15) from the separator operating at 1.6 MPa is then sent to a stabilization column (16) operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds of the liquid fraction as well as the hhS (line 17) and a bottom section freed from H ₂ S, called stabilized gasoline (line 18).

The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in Table 14.

Table 14: Characteristics of the gaseous fraction (line 14) after gas/liquid separation of the reaction effluent 12

The characteristics of the gasoline cut obtained after stabilization (line 18) are illustrated in Table 15.

Table 15: Characteristics of heavy gasoline after the first and second hydrodesulphurization stages The loss of C5+ compounds in the gaseous fraction obtained after the separation stage (line 14) and in the gaseous cut after stabilization of the gasoline (line 17) is given in Table 16.

Table 16: Loss of C5+ compounds between the gasoline obtained after the second hydrodesulfurization step (line 12) and the gasoline obtained after the stabilization step (line 18)

The process according to Example 5 makes it possible to obtain a heavy gasoline with a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 46.6% by weight (relative). This octane loss is greater than in the example according to the invention.

All of the results are compared in Table 17 below, presenting the olefin contents in the gasoline and losses of C5+ compounds in the gaseous purges after the stabilization step. Carrying out the process according to the invention appears to be the best compromise for maximizing the content of olefins in the gasoline obtained on the one hand, and for limiting the loss of recoverable C5+ compounds in a gasoline cut.

Table 17

Claims

34 CLAIMS

1. Process for treating a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps: a) contacting in at least one reactor, gasoline, hydrogen and a hydrodesulphurization catalyst comprising an oxide support and an active phase comprising a group VIB metal and a group VIII metal, at a temperature of between 210 and 320°C, at a pressure of between 1.5 and 3 MPa , with a space velocity comprised between 1 and 10 h ^-1 and a ratio between the hydrogen flow rate expressed in normal m ³ per hour and the feed rate to be treated expressed in m ³ per hour at standard conditions comprised between 100 and 600 Nm ³ /m ³ , so as to convert at least part of the sulfur compounds into H2S; b) in at least one reactor, the effluent resulting from stage a) without elimination of the H2S formed, is brought into contact with hydrogen, and a hydrodesulphurization catalyst comprising an oxide support and a phase active consisting of at least one group VIII metal, at a temperature of between 280 and 400°C, at a pressure of between 1.0 and 3 MPa, with a space velocity of between 1 and 10 ^h'1 and a ratio between the hydrogen flow rate expressed in normal m ³ per hour and the feed flow rate to be treated expressed in m ³ per hour at standard conditions of between 100 and 600 Nm ³ /m ³ ; c) the effluent from step b) is sent to at least one separation drum operating at a pressure of between 1.0 and 2.0 MPa to obtain a gaseous fraction containing H2S and hydrogen and a liquid fraction containing the desulfurized gasoline and a fraction of residual H2S; d) said liquid fraction obtained in step c) is sent to a stabilization column to obtain at the top a stream comprising the residual H ₂ S and C4- hydrocarbon compounds and at the bottom a stabilized gasoline; e) the gaseous fraction obtained at the end of step c) is recycled at least in part to at least one of steps a) and/or b).

2. Method according to claim 1, in which the pressure of the separation drum of step c) is between 1.2 and 1.8 MPa.

3. Process according to any one of claims 1 and 2, in which the catalyst of step a) comprises a group VIII metal content of between 0.1 and 10% by weight of oxide of the group VIII metal per relative to the total weight of the catalyst, and a group VIB metal content of between 1 and 20% by weight of group VIB metal oxide relative to the total weight of the catalyst. 35

4. Process according to any one of claims 1 to 3, in which the catalyst of step a) comprises a molar ratio of metal from group VIII to metal from group VI B of the catalyst of between 0.1 and 0.8.

5. Process according to any one of claims 1 to 4, in which the catalyst of step a) comprises a specific surface of between 5 and 400 m ² /g.

6. Process according to any one of claims 1 to 5, in which the catalyst of step a) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content of weight relative to the total weight of cobalt oxide catalyst, in CoO form, of between 0.1 and 10%, a content by weight relative to the total weight of molybdenum oxide catalyst, in MoOa form, of between 1 and 20%, a cobalt/molybdenum molar ratio of between 0.1 and 0.8, a content by weight relative to the total weight of catalyst of phosphorus oxide in the P2O5 form of between 0.3 and 10% when the phosphorus is present, said catalyst having a specific surface between 50 and 250 m ² /g.

7. Process according to any one of claims 1 to 6, in which the catalyst of step b) comprises a group VIII metal content of between 1 and 60% by weight of oxide of the group VIII metal relative to the total catalyst weight.

8. Process according to any one of claims 1 to 7, in which the catalyst of step b) comprises a specific surface of between 5 and 400 m ² /g.

9. Process according to any one of claims 1 to 8, in which the catalyst of step b) consists of alumina and nickel, said catalyst containing a content by weight relative to the total weight of oxide catalyst nickel, in NiO form, between 5 and 20%, said catalyst having a specific surface between 30 and 180 m ² /g.

10. Process according to any one of claims 1 to 9, in which the temperature of step b) is higher than the temperature of step a).

11. Process according to claim 10, in which the temperature of step b) is higher by at least 5°C than the temperature of step a).

12. Method according to any one of claims 1 to 11, in which before step a) a gasoline distillation step is carried out so as to split said gasoline into at least two light and heavy gasoline cuts and the heavy gasoline cut in steps a), b), c), d) and e).

13. Process according to any one of claims 1 to 12, in which before step a) and before any possible distillation step, the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst to selectively hydrogenating the diolefins contained in said gasoline into olefins.

14. Process according to any one of claims 1 to 13, in which the gasoline is a catalytic cracked gasoline.