Process for producing a low-sulfur light gasoline cut

Abstract

A process for producing a light gasoline comprising a sulfur content of less than 10 ppm by weight relative to the total weight of said light gasoline from a gasoline containing sulfur compounds, olefins and diolefins, which process comprises: a) a selective hydrogenation step so as to hydrogenate the diolefins and carry out a reaction to make a portion of the sulfur compounds heavier; b) a step of separating the effluent obtained at the end of step a) into a gaseous fraction, a light gasoline cut and a heavy gasoline cut, said step b) being carried out in a fractionation column comprising n trays, n being an integer greater than or equal to 20, the first tray being the reboiler and tray "n" being the condenser, it being understood that the light gasoline cut is withdrawn from said fractionation column at tray "ni", with i between 1 and 10.

Description

Title of the invention: Process for producing a light gasoline cut with low sulfur content Field of invention

[0001] The present invention relates to a process for reducing the content of sulfur compounds in an olefinic gasoline, so as to produce a so-called desulfurized gasoline, while limiting the loss of octane induced by the hydrogenation of olefins. State of the art

[0002] The production of gasolines meeting new environmental standards requires a significant reduction in their sulfur content.

[0003] It is also known that conversion gasolines, and more particularly those from fluidized bed catalytic cracking, which can represent 30 to 50% of the gasoline pool, have high mono-olefin and sulfur contents.

[0004] The sulfur present in gasolines is for this reason attributable, at nearly 90%, to gasolines resulting from fluidized bed catalytic cracking processes, which will be called in the following FCC gasolines (Fluid Catalytic Cracking according to the English terminology, which can be translated as fluidized bed catalytic cracking). FCC gasolines therefore constitute the preferred feedstock for the process of the present invention.

[0005] Among the possible ways to produce low-sulfur fuels, the one that is very widely adopted consists of specifically treating sulfur-rich gasoline bases by catalytic hydrodesulfurization processes in the presence of hydrogen. Traditional processes desulfurize gasolines non-selectively by hydrogenating a large portion of the mono-olefins, which results in a high loss in octane number and high hydrogen consumption. The most recent processes, such as the Prime G+ process (trademark), make it possible to desulfurize olefin-rich cracked gasolines, while limiting the hydrogenation of the mono-olefins, and consequently the resulting octane loss and high hydrogen consumption. Such processes are for example described in patent applications EP1077247 and EPI 174485.

[0006] The residual sulfur compounds generally present in desulfurized gasoline can be separated into two distinct families: the unconverted refractory sulfur compounds present in the feedstock on the one hand, and the sulfur compounds formed in the reactor by secondary reactions called recombination. 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.

[0007] Mercaptans, with the chemical formula R-SH, where R is an alkyl group, are also called recombinant mercaptans. Their formation or decomposition follows the thermodynamic equilibrium of the reaction between monoolefins and hydrogen sulfide to form recombinant mercaptans. An example is illustrated according to the following reaction:

[0008] [Chem.l] A + H2S 5........................> Y SH

[0009] The sulfur contained in recombinant mercaptans generally represents between 20% and 80% by weight of the residual sulfur in desulfurized gasolines.

[0010] The formation of recombinant mercaptans is in particular described in patent US6231754 and application WO01 / 40409 which teach various combinations of operating conditions and catalysts making it possible to limit the formation of recombinant mercaptans.

[0011] Other solutions to the problem of the formation of recombinant mercaptans are based on a treatment of partially desulfurized gasolines to extract said recombinant mercaptans. Some of these solutions are described in applications WO02 / 28988 or WO01 / 79391.

[0012] Still other solutions are described in the literature for desulfurizing FCC gasolines using a combination of hydrodesulfurization steps and elimination of recombinant mercaptans by reaction into thioethers or disulfides (also called softening or "sweetening" according to English terminology) (see for example US7799210, US6960291, US2007 / 114156, EP2861094).

[0013] Document US2018 / 0171244 discloses a method for treating gasoline comprising a step of separating the feed into a light gasoline cut and a heavy gasoline cut, in which the light gasoline cut is sent to a hydrodesulfurization unit in order to reduce the presence of sulfur in said light fraction. It is therefore necessary to treat the light gasoline cut after separation in order to limit the quantity of total sulfur contained in said light gasoline cut.

[0014] An aim of the present invention is to provide a method for treating a gasoline containing sulfur compounds, olefins and diolefins to directly obtain a light gasoline with a low sulfur content and a low mercaptan content, while limiting the loss of octane number, and therefore being able to be sent directly to the gasoline pool for its use as fuel, without undergoing additional treatment. Objects of the invention

[0015] The subject of the present invention is a process for producing a light gasoline (here also called LCN gasoline or "Light Cracked Naphtha" according to English terminology) comprising a sulfur content of less than 10 ppm by weight relative to the total weight of said light gasoline from a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps:

[0016] a) a selective hydrogenation step so as to hydrogenate the diolefins and carry out a reaction to make a portion of the sulfur compounds heavier, in which process the gasoline and hydrogen are brought into contact with a selective hydrogenation catalyst, at a temperature of between 100°C and 220°C, with a liquid space velocity of between 1 h 1 and 7 h 1 and a pressure of between 0.5 MPa and 5 MPa, and with a molar ratio between hydrogen and the diolefins to be hydrogenated of greater than 1 mol / mol and less than 100 mol / mol, with a ratio between the hydrogen flow rate expressed in normal m3 per hour and the feed flow rate to be treated expressed in m3 per hour under standard conditions of between 2 Nm3 / m3 and 100 Nm3 / m3, said selective hydrogenation catalyst comprising an active phase containing at least one metal from group VIB and at least one metal from group VIB. at least one metal from group VIII, and a porous support containing at least alumina,said catalyst comprising a specific surface area of ​​between 100 m2 / g and 400 m2 / g;

[0017] b) a step of separating the effluent obtained at the end of step a) into a gaseous fraction, a light gasoline cut (LCN) and a heavy gasoline cut (also called here HCN or “heavy cracked naphtha” according to English terminology), said step b) being carried out in a fractionation column comprising n trays, n being an integer greater than or equal to 20, the first tray being the reboiler and tray “n” being the condenser, it being understood that the light gasoline cut (LCN) is withdrawn from said fractionation column at tray “ni”, with i between 1 and 10.

[0018] The Applicant has surprisingly discovered that it is possible to directly obtain a light gasoline cut meeting the required specifications and enriched in olefins from a gasoline feedstock containing sulfur compounds, olefins and diolefins by carrying out a step of selective hydrogenation of said gasoline feedstock under specific operating conditions and in the presence of a specific catalyst, then a step of separation of the effluent obtained at the end of the selective hydrogenation step, the separation being carried out in a fractionation column with withdrawal of the light gasoline cut at a well-defined level of the column.

[0019] According to one or more embodiments, said fractionation column comprises between 20 and 100 trays.

[0020] According to one or more embodiments, said light gasoline cut is withdrawn from said fractionation column at the “ni” tray with i between 1 and 6.

[0021] According to one or more embodiments, said metal from group VIII is chosen to be nickel.

[0022] According to one or more embodiments, said group VIB metal is molybdenum.

[0023] According to one or more embodiments, step a) is carried out in the presence of a catalyst comprising nickel with a content by weight of nickel oxide, in NiO form, of between 1% and 12%, and molybdenum with 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.

[0024] According to one or more embodiments, said catalyst comprises a specific surface area of ​​between 100 m2 / g and 280 m2 / g.

[0025] According to one or more embodiments, the method further comprises a step c) of hydrodesulfurization of the heavy gasoline HCN cut obtained at the end of step b) in the presence of hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase comprising a metal from group VIB and a metal from group VIII, at a temperature of between 210°C and 320°C, at a pressure of between 1 MPa and 4 MPa, with a space velocity of between 1 h 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of feed to be treated expressed in m3 per hour at standard conditions of between 100 Nm3 / m3 and 600 Nm3 / m3, so as to convert at least a portion of the sulfur compounds into H2S.

[0026] According to one or more embodiments, the hydrodesulfurization catalyst of step c) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said hydrodesulfurization 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%, 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 P2O5 form of between 0.3 and 10% when phosphorus is present, said hydrodesulfurization catalyst having a specific surface area of ​​between 50 m2 / g and 250 m2 / g.

[0027] According to one or more embodiments, the method further comprises a step d) of finishing hydrodesulfurization of the effluent obtained from step c) without elimination of the H2S formed, in the presence of hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, at a temperature between 280°C and 400°C, at a pressure between 0.5 MPa and 5 MPa, with a space velocity between 1 h 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of charge to be treated expressed in m3 per hour under standard conditions between 100 Nm3 / m3 and 600 Nm3 / m3;

[0028] According to one or more embodiments, the hydrodesulfurization catalyst of step d) consists of alumina and nickel, said hydrodesulfurization catalyst containing a content by weight, relative to the total weight of catalyst, of nickel oxide, in NiO form, of between 5 and 20%, said hydrodesulfurization catalyst having a specific surface area of ​​between 30 m2 / g and 180 m2 / g.

[0029] According to one or more embodiments, the method further comprises a step e) of separating the H2S formed and present in the effluent from step d).

[0030] According to one or more embodiments, the gasoline is a catalytic cracking gasoline. List of figures

[0031] [Fig-1] [Fig.l] is a schematic representation of the method according to the invention.

[0032] [Fig.2] [Fig.2] is a schematic representation of a process not in accordance with the invention. Detailed description Definitions

[0033] 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 IUP AC classification.

[0034] Specific surface area means the BET specific surface area (SBet in m2 / g) determined by nitrogen adsorption in accordance with ASTM D 3663-78 established from the BRUNAUER-EMMETT-TELLER method described in the periodical "The Journal of American Society", 1938, 60, 309.

[0035] The total pore volume of the catalyst or support used for the preparation of the catalyst is understood to mean 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.

[0036] The wetting angle was taken equal to 140° following the recommendations of the work “Techniques de l'ingénieur, traité analyse et caractérisation”, pages 1050-1055, written by Jean Charpin and Bernard Rasneur. In order to obtain a better To be precise, the total pore volume value is the total pore volume value measured by mercury porosimeter intrusion measured on the sample minus the total pore volume value measured by mercury porosimeter intrusion measured on the same sample at a pressure corresponding to 30 psi (approximately 0.2 MPa).

[0037] The contents of elements of group VIII, group VIB and phosphorus are measured by X-ray fluorescence. Process

[0038] The present invention relates to a process for producing a light gasoline comprising a sulfur content of less than 10 ppm by weight relative to the total weight of said light gasoline from a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps:

[0039] a) a selective hydrogenation step so as to hydrogenate the diolefins and carry out a reaction to make a portion of the sulfur compounds heavier, in which process the gasoline and hydrogen are brought into contact with a selective hydrogenation catalyst, at a temperature of between 100°C and 220°C, with a liquid space velocity of between 1 h 1 and 7 h 1 and a pressure of between 0.5 MPa and 5 MPa, and with a molar ratio between hydrogen and the diolefins to be hydrogenated of greater than 1 mol / mol and less than 100 mol / mol,with a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of charge to be treated expressed in m3 per hour under standard conditions of between 2 Nm3 / m3 and 100 Nm3 / m3, said selective hydrogenation catalyst comprising an active phase containing at least one metal from group VIB and at least one metal from group VIII, and a porous support containing at least alumina, said catalyst comprising a specific surface area of ​​between 100 m2 / g and 400 m2 / g; ,

[0040] b) a step of separating the effluent obtained at the end of step a) into a gaseous fraction, a light gasoline cut and a heavy gasoline cut, said step b) being carried out in a fractionation column comprising “n” trays, n being an integer greater than or equal to 20, the first tray being the reboiler (i.e. at the bottom of the fractionation column) and tray “n” being the condenser (i.e. at the top of the fractionation column), it being understood that the light gasoline cut is withdrawn from said fractionation column at tray “ni” with i between 1 and 10;

[0041] c) optionally, a step of hydrodesulfurization of the heavy gasoline fraction obtained at the end of step b) in the presence of hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase comprising a metal from group VIB and a metal from group VIII, at a temperature between 210°C and 320°C, at a pressure between 1 MPa and 4 MPa, with a space velocity between 1 h 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of the load to be treated expressed in m3 per hour under standard conditions between 100 Nm3 / m3 and 600 Nm3 / m3, so as to convert at least part of the sulfur compounds into H2S;

[0042] d) optionally, a finishing hydrodesulfurization step of the effluent obtained from step c) without elimination of the H2S formed, in the presence of hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, at a temperature between 280°C and 400°C, at a pressure between 0.5 MPa and 5 MPa, with a space velocity between 1 h 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the feed flow rate to be treated expressed in m3 per hour at standard conditions between 100 Nm3 / m3 and 600 Nm3 / m3;

[0043] e) optionally, a step of separation of the H2S formed and present in the effluent from step d). Description of the load

[0044] The process according to the invention makes it possible to treat any type of gasoline cut containing sulfur compounds, olefins and diolefins, such as for example a cut from a coking unit (coking according to English terminology), vis-coreduction (visbreaking according to English terminology), steam cracking (steam cracking according to English terminology) or fluidized bed catalytic cracking (FCC, Fluid Catalytic Cracking according to English 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 English terminology)) or from conversion processes (coking or steam cracking gasoline). Said charge preferably consists of a gasoline cut from a catalytic cracking unit.

[0045] The feedstock is a gasoline cut containing sulfur compounds and olefins whose boiling point range typically extends from the boiling points of hydrocarbons with 2 or 3 carbon atoms (C2 or C3) up to 260°C, preferably from the boiling points of hydrocarbons with 2 or 3 carbon atoms (C2 or C3) up to 220°C, more preferably from the boiling points of hydrocarbons with 5 carbon atoms up to 220°C. The process according to the invention can also treat feedstocks having end points lower than those mentioned above, such as for example a C5-180°C cut.

[0046] The sulfur content of gasoline cuts produced by FCC depends on the sulfur content of the feedstock treated by the FCC, the presence or absence of pretreatment of the FCC feedstock, and the end point of the cut. Generally, the sulfur contents of the entire gasoline cut, especially if it comes from the FCC, are greater than 100 ppm by weight, and most of the time greater than 500 ppm by weight. For gasolines with end points above 200°C, sulfur contents are often greater than 1000 ppm by weight, and in some cases can even reach values ​​of the order of 4000 ppm to 5000 ppm by weight.

[0047] The feedstock treated by the process according to the invention may be a feedstock containing sulfur compounds in a content greater than 1000 ppm by weight of sulfur, and often greater than 1500 ppm.

[0048] Furthermore, gasolines from FCC units 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% by weight of sulfur, of which generally less than 300 ppm of mercaptans. Step a): Selective hydrogenation

[0049] The selective hydrogenation according to step a) of the process according to the invention consists mainly of:

[0050] - selectively hydrogenate diolefins into monoolefins;

[0051] - transform saturated light sulfur compounds and mainly mercaptans, into heavier sulfides or mercaptans by reaction with mono-olefins;

[0052] - isomerize mono-olefin compounds having their external C=C double bond into their internal C=C double bond isomer.

[0053] The hydrogenation reactions of diolefins to monoolefins are illustrated below by the transformation of 1,3 pentadiene, an unstable compound, which can easily be hydrogenated to pent-2-ene. However, we seek to limit the secondary hydrogenation reactions of monoolefins which in the example below would lead to the formation of n-pentane.

[0054] [Chem.2] CH «2 ch„ CH H2 CHU CH2 ""CH bH3 ------H / L

[0055] The sulfur compounds that are sought to be transformed are mainly mercaptans. The main reaction for transforming mercaptans consists of a thioetherification reaction between monoolefins and mercaptans. This reaction is illustrated below by the addition of propane-2-thiol to pent-2-ene to form a propyl-pentyl sulfide.

[0056] [Chem.3]

[0057] In the presence of hydrogen, the transformation of sulfur compounds can also occur through the intermediate formation of H2S which can then be added to the unsaturated compounds present in the feed. This pathway is, however, a minority in the preferred reaction conditions.

[0058] In addition to mercaptans, the compounds likely to be transformed and made heavier in this way are sulfides and mainly CS2, COS, thiophane, methyl-thiophane.

[0059] In certain cases, reactions of weighting of light nitrogen compounds, and mainly nitriles, pyrrole and its derivatives, can also be observed.

[0060] According to the invention, the catalyst also makes it possible to carry out an isomerization of mono-olefinic compounds having their C=C double bond in the external position into their isomer having their C=C double bond in the internal position.

[0061] This reaction is illustrated below by the isomerization of hexene-1 into hexene-2 ​​or hexene-3.

[0062] [Chem.4] tLC ^CH3

[0063] In the selective hydrogenation process according to the invention, the feedstock to be treated is mixed with hydrogen before being brought into contact with the catalyst. The quantity of hydrogen injected is such that the molar ratio between hydrogen and the diolefins to be hydrogenated is greater than 1 (stoichiometry) and less than 100, and preferably between 1 and 10 mol / mol. Too large an excess of hydrogen can lead to strong hydrogenation of the monoolefins and consequently, a reduction in the octane number of the gasoline.

[0064] The entire feedstock is generally injected at the reactor inlet. However, it may be advantageous, in certain cases, to inject a fraction or the entire feedstock between two consecutive catalytic beds placed in the reactor. This embodiment makes it possible in particular to continue operating the reactor if the reactor inlet becomes blocked by deposits of polymers, particles, or gums present in the feedstock.

[0065] The mixture consisting of gasoline and hydrogen is brought into contact with the catalyst at a temperature between 100°C and 220°C, and preferably between 110°C and 200°C, with a liquid space velocity (WH) between 1 h 1 and 7 h1, the unit of liquid space velocity being the volume in m3 per hour at standard conditions, per m3 of catalyst. The pressure is adjusted so that the reaction mixture is predominantly in liquid form in the reactor. The pressure is between 0.5 MPa and 5 MPa and preferably between 1 MPa and 4 MPa.

[0066] The reaction of selective hydrogenation of diolefins and weighting of light mercaptans is preferably carried out on a catalyst comprising at least one metal from group VIII and at least one metal from group VIB and a porous support containing at least alumina.

[0067] The group VIII metal content of the active phase, measured in oxide form, is between 1 and 20% by weight relative to the total weight of the catalyst, preferably between 2 and 15% by weight, and even more preferably between 4 and 13% by weight. The group VIII metal is preferably chosen from nickel, cobalt, and iron. More preferably, the group VIII metal is nickel.

[0068] The group VIB metal content of the active phase, measured in oxide form, is between 1 and 18% by weight relative to the total weight of the catalyst, preferably between 1 and 15% by weight, and even more preferably between 2 and 13% by weight. The group VIB metal is preferably chosen from molybdenum and tungsten. More preferably, the group VIB metal is molybdenum.

[0069] The element of group VIII is preferably chosen from nickel and cobalt and in particular nickel. The element of group VIB is preferably chosen from molybdenum and tungsten and preferably molybdenum.

[0070] According to the invention, the catalyst used in the selective hydrogenation step has a specific surface area of ​​between 100 m2 / g and 400 m2 / g, preferably of between 100 m2 / g and 300 m2 / g, preferably of between 100 m2 / g and 280 m2 / g.

[0071] The specific surface area 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”, Academie Press, 1999, for example using an Autopore III™ model device from the Microméritics™ brand.

[0072] The pore volume of the selective hydrogenation catalyst is generally between 0.4 cm3 / g and 1.3 cm3 / g, preferably between 0.5 cm3 / g and 1.1 cm3 / g.

[0073] The total pore volume is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140°, as described in the same work.

[0074] 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, said catalyst having a specific surface area of ​​between 100 m2 / g and 400 m2 / g, preferably between 100 m2 / g and 300 m2 / g, preferably between 100 m2 / g and 280 m2 / g. The sulfurization rate of the metals constituting the catalyst is, preferably, greater than 60%.

[0075] After selective hydrogenation, the diolefin content, determined by means of the maleic anhydride index (MAV or "Maleic Anhydride Value" according to English terminology), according to the UOP 326 method, is generally reduced to less than 6 mg maleic anhydride / g (MA / g), or even less than 4 mg MA / g and more preferably less than 2 mg MA / g. In certain cases, less than 1 mg MA / g can be obtained. Step b) Separation

[0076] The effluent obtained at the end of step a) is then sent to a fractionation column (also called a splitter in English terminology) to obtain a gaseous fraction, a light gasoline cut (LCN) and a heavy gasoline cut (HCN).

[0077] According to an essential aspect of the invention, step b) is carried out in a fractionation column comprising “n” trays, n being an integer greater than or equal to 20, the first tray being the reboiler and tray “n” being the condenser, it being understood that the light gasoline is withdrawn from said fractionation column at tray “ni”, with i between 1 and 10, preferably between 2 and 6, and more preferably between 3 and 6. The number “n” of trays in the fractionation column is preferably between 20 and 100, preferably between 20 and 60.

[0078] The fractionation column generally operates at a pressure between 0.1 MPa and 2 MPa, and preferably between 0.2 MPa and 1 MPa.

[0079] A hydrogen top-up can be carried out in the column in order to maintain the pressure in the event of total hydrogen consumption during step a) of the process according to the invention.

[0080] Advantageously, the fractionation column comprises a reboiler whose reboiling steam is at a temperature below 260°C in order to limit the subsequent cracking (back cracking according to English terminology) of heavy mercaptans into light mercaptans in the heavy gasoline cut.

[0081] The reflux ratio, expressed as the ratio of the liquid traffic in the column divided by the distillate flow rate expressed in kg / h, is generally less than 1, preferably less than 0.9.

[0082] The gaseous fraction is obtained by condensation of the gas phase produced at the top of the fractionation column then by gas / liquid separation which allows the elimination of the hydrogen present in the effluent obtained at the end of step a) and the production of a liquid phase consisting of the C5 to C7 hydrocarbon compounds entrained at the top of the column, which is returned as reflux to the column. Condensation is generally carried out by cooling to a temperature between 40°C and 65°C. The liquid reflux makes it possible to control the temperature at the top of the column, thus enabling the control of the sulfur content of the light fraction of the gasoline.

[0083] Thus preferably:

[0084] - the gaseous fraction mainly comprises hydrogen, possibly hydrocarbon compounds Cl to C4 which may come from the hydrogen make-up gas flow;

[0085] - the LCN light gasoline cut is a C5 hydrocarbon cut (i.e. containing hydrocarbons having 5 or less than 5 carbon atoms per molecule);

[0086] - the heavy gasoline HCN cut is a C6+ cut (i.e. containing hydrocarbons can have 6 or more carbon atoms per molecule);

[0087] Advantageously, the total sulfur content of the light gasoline is less than 10 ppm by weight. The light gasoline obtained does not require an additional hydrodesulfurization step and can be sent directly to the gasoline pool.

[0088] Step c) Hydrodesulfurization of the heavy gasoline fraction (HCN) [optional]

[0089] The hydrodesulfurization step c) can be implemented to reduce the sulfur content of the heavy gasoline (HCN) to be treated by converting the sulfur compounds into H2 S which can then be eliminated in step e).

[0090] The hydrodesulfurization step c) consists of bringing the heavy 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.

[0091] According to a preferred embodiment of the invention, step c) is carried out with the aim of carrying out hydrodesulfurization selectively, i.e. with a hydrogenation rate of the mono-olefins of less than 80%, preferably less than 70% and very preferably less than 60%.

[0092] The temperature is generally between 210°C and 320°C and preferably between 220°C 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 the hydrodesulfurization step c) is carried out in several reactors in series, the temperature of each reactor is generally at least 5°C higher, preferably at least 10°C higher and very preferably at least 30°C higher than the temperature of the reactor preceding it.

[0093] The operating pressure of this step is generally between 1 MPa and 4 MPa and preferably between 1.5 MPa and 3 MPa.

[0094] 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 m3 per hour at standard conditions, per m3 of catalyst (also called space velocity) is between 1 h 1 and 10 h 1 and preferably between 2 h 1 and 8 h

[0095] The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m3 per hour (Nm3 / h) and the flow rate of feedstock to be treated expressed in m3 per hour at standard conditions (15°C, 0.1 MPa) is between 100 Nm3 / m3 and 600 Nm3 / m3, preferably between 200 Nm3 / m3 and 500 Nm3 / m3. Normal m3 means the quantity of gas in a volume of 1 m3 at 0°C and 0.1 MPa.

[0096] The hydrogen required for this step may be fresh hydrogen or recycled hydrogen, preferably freed from H2S, or a mixture of fresh hydrogen and recycled hydrogen. Preferably, fresh hydrogen will be used.

[0097] The desulfurization rate of step c), which depends on the sulfur content of the feed to be treated, is generally greater than 50% and preferably greater than 70% so that the product from step c) contains less than 100 ppm by weight of sulfur and preferably less than 50 ppm by weight of sulfur.

[0098] The catalyst used in step c) must have good selectivity with respect to hydrodesulfurization reactions compared to the olefin hydrogenation reaction.

[0099] The hydrodesulfurization catalyst of step c) comprises an oxide support and an active phase comprising a group VIB metal and a group VIII metal and optionally phosphorus and / or an organic compound as described below.

[0100] 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 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.

[0101] The content of group VIII metal 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.

[0102] The content of group VIB metal 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% by weight.

[0103] 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.

[0104] Optionally, the catalyst may also have a phosphorus content generally between 0.3 and 10% by weight of P2O5 relative to the total weight of ca talyser, preferably between 0.3 and 5% by weight, very preferably between 0.5 and 3% by weight. For example, the phosphorus present in the catalyst is combined with the metal of group VIB and optionally also with the metal of group VIII in the form of heteropolyanions.

[0105] 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.

[0106] Preferably, the catalyst is characterized by a specific surface area of ​​between 5 m2 / g and 400 m2 / g, preferably between 10 m2 / g and 250 m2 / g, preferably between 50 m2 / g and 250 m2 / g. The specific surface area 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”, Academie Press, 1999, for example using an Autopore III™ model device from the Micro-méritics™ brand.

[0107] The total pore volume of the catalyst is generally between 0.4 cmVg and 1.3 cmVg, preferably between 0.6 cmVg and 1.1 cmVg. The total pore volume is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140°, as described in the same work.

[0108] The packed filling density (DRT) of the catalyst is generally between 0.4 g / mL and 0.8 g / mL, preferably between 0.4 g / mL and 0.7 g / mL. The DRT measurement consists of introducing the catalyst into a test tube whose volume has been previously determined and then, by vibration, packing 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.

[0109] The catalyst can be in the form of small diameter extrudates, cylindrical or multi-lobed (tri-lobed, quadri-lobed, etc.), or spheres.

[0110] The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas 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 silica-aluminas, very preferably, the oxide support is essentially made 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 is preferably made solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, i.e. which contains aluminas of theta, delta, kappa or alpha phase, alone or in a mixture and a quantity of less than 20% of alumina. gamma, chi or eta phase.

[0111] The catalyst may also further comprise at least one organic compound containing oxygen and / or nitrogen and / or sulfur before sulfurization.

[0112] A very preferred embodiment of the invention corresponds to the use for step c) 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 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%, 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 P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface area of ​​between 50 m2 / g and 250 m2 / 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..

[0113] Step d): Finishing hydrodesulfurization [optional]

[0114] During the hydrodesulfurization step d), a large part of the sulfur compounds is transformed 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 c) to the mono-olefins present in the feedstock.

[0115] This so-called “finishing” hydrodesulfurization step is mainly implemented to reduce the content of recombinant mercaptans. Preferably, step d) is carried out at a higher temperature than that of step c). Indeed, by using a higher temperature in this step compared to the temperature of step c), the formation of olefins and H2S will be favored by the thermodynamic equilibrium. Step d) also makes it possible to hydrodesulfurize the more refractory sulfur compounds.

[0116] The hydrodesulfurization step d) consists of bringing the effluent from step c) into contact, optionally with the addition of hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization.

[0117] The hydrodesulfurization step d) is carried out without significant hydrogenation of the olefins. The hydrogenation rate of the olefins of the catalyst of the hydrodesulfurization step d) is generally less than 5%, and even more generally less than 2%.

[0118] The temperature of this step is generally between 280°C and 400°C, more preferably between 290°C and 380°C, and very preferably between 300°C and 360°C. The temperature of this step d) is generally at least 5°C higher, preferably at least 10°C and very preferably at least 30°C at the temperature of step c).

[0119] The operating pressure of this step is generally between 0.5 MPa and 5 MPa and preferably between 1 MPa and 3 MPa.

[0120] 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 m3 per hour at standard conditions, per m3 of catalyst (also called space velocity) is between 1 h 1 and 10 h 1 and preferably between 2 h 1 and 8 h

[0121] Preferably, the hydrogen flow rate is constant and equal to the quantity injected in step c) less the hydrogen consumed in step c). The hydrogen flow rate is generally such that the ratio between the hydrogen flow rate expressed in normal m3 per hour (Nm3 / h) and the flow rate of feedstock to be treated expressed in m3 per hour at standard conditions (15°C, 0.1 MPa) is between 100 Nm3 / m3 and 600 Nm3 / m3, preferably between 200 Nm3 / m3 and 500 Nm3 / m3.

[0122] The desulfurization rate of step d), 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 from step d) contains less than 60 ppm by weight of sulfur and preferably less than 40 ppm by weight of sulfur, and very preferably less than 20 ppm by weight of sulfur.

[0123] The hydrodesulfurization steps c) and d) can be carried out either in a single reactor containing the two catalysts, or in at least two different reactors. When steps c) and d) are carried out using two reactors, the latter two are placed in series, the second reactor treating all of the effluent at the outlet of the first reactor (without separation of the liquid and gas between the first and second reactors).

[0124] The catalyst of step d) is of a different nature and / or composition from that used in step c). The catalyst of step d) is in particular a very selective hydrodesulfurization catalyst: it makes it possible to hydrodesulfurize without hydrogenating the olefins and therefore to maintain the octane number.

[0125] The catalyst which may be suitable for this step d) 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 metal from group VIII, 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 metal from group VIII, preferably nickel. Particularly preferably, the active phase consists of nickel.

[0126] The content of group VIII metal is between 1 and 60% by weight of group VIII metal oxide relative to the total weight of the catalyst, preferably between between 5 and 30% by weight, very preferably between 5 and 20% by weight.

[0127] Preferably, the catalyst is characterized by a specific surface area of ​​between 5 m2 / g and 400 m2 / g, preferably between 10 m2 / g and 250 m2 / g, preferably between 20 and 200 m2 / g, very preferably between 30 m2 / g and 180 m2 / g. The specific surface area 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”, Academie Press, 1999, for example using an Autopore III™ model device from the Microméritics™ brand.

[0128] The pore volume of the catalyst is generally between 0.4 cmVg and 1.3 cm3 / g, preferably between 0.6 cmVg and 1.1 cmVg. The total pore volume is measured by mercury porosimetry according to ASTM D4284 with a wetting angle of 140°, as described in the same work.

[0129] The tapped filling density (TFD) of the catalyst is generally between 0.4 g / mL and 0.8 g / mL, preferably between 0.4 g / mL and 0.7 g / mL.

[0130] The DRT measurement consists of introducing the catalyst into a test tube whose volume has been previously determined and then, by vibration, tamping it until a constant volume is obtained. The apparent density of the tamped product is calculated by comparing the mass introduced and the volume occupied after tamping.

[0131] The catalyst can be in the form of small diameter extrudates, cylindrical or multi-lobed (tri-lobed, quadri-lobed, etc.), or spheres.

[0132] The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas 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 silica-aluminas, very preferably, the oxide support is essentially made 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 is preferably made solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, i.e. which contains aluminas of theta, delta, kappa or alpha phase, alone or in a mixture and a quantity of less than 20% of alumina of gamma, chi or eta phase.

[0133] A very preferred embodiment of the invention corresponds to the use for step d) of a catalyst consisting of alumina and 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 20%, said catalyst having a specific surface area of ​​between 30 m2 / g and 180 m2 / g.

[0134] The catalyst of the hydrodesulfurization step d) is characterized by a catalytic hydrodesulfurization activity generally between 1% and 90%, preferably between 1% and 70%, and very preferably between 1% and 50% of the catalytic activity of the catalyst of the hydrodesulfurization step c). Step e): Separation of H2S [optional]

[0135] This step is implemented in order to separate the excess hydrogen as well as the H2S formed during steps c) and d). Any method known to those skilled in the art can be considered.

[0136] According to a first embodiment, after the hydrodesulfurization steps c) and d), the effluent is cooled to a temperature generally lower than 80°C in order to condense the hydrocarbons. The gas and liquid phases are then separated in a separation tank. The liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved H2S is sent to a stabilization column or debutanizer. This column separates a top cut essentially consisting of residual H2S and hydrocarbon compounds having a boiling point lower than or equal to that of butane and a bottom cut freed from H2S, called stabilized gasoline, containing the compounds having a boiling point higher than that of butane.

[0137] According to a second embodiment, after the condensation step, the liquid fraction which contains the desulfurized gasoline as well as a fraction of the dissolved 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 water vapor, in a distillation column in order to extract at the top, the light compounds which have been entrained by dissolution in the liquid fraction as well as the dissolved residual H2S. The temperature of the stripped gasoline recovered at the bottom of the column is generally between 120°C and 250°C.

[0138] Preferably, the separation step e) is carried out in a stabilization column or debutanizer. Indeed, a stabilization column makes it possible to separate H2S more efficiently than a stripping section.

[0139] Step e) is preferably carried out so that the sulfur in the form of H2S remaining in the desulfurized gasoline represents less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.

[0140] Description of the preparation of catalysts and sulfurization

[0141] The preparation of the catalysts of steps c) and d) is known and generally comprises a step of impregnation of the metals of group VIII and group VIB when present, and optionally of the phosphorus and / or the organic compound on the oxide support, followed by drying and then optional calcination to obtain the active phase in their oxide forms. Before 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.

[0142] 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 solubilize the metal precursors in the desired concentrations.

[0143] For example, among the sources of molybdenum, it is possible to use oxides and hydroxides, molybdic acids and their salts, in particular ammonium salts such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PMoi204o), and their salts, and possibly silicomolybdic acid (H4 SiMoi204o) and its salts. The sources of molybdenum can also be any heteropolycompound of the Keggin, lacunar Keggin, substituted Keggin, Dawson, Anderson, Strandberg type, for example. Molybdenum trioxide and heteropolycompounds of the Keggin, lacunar Keggin, substituted Keggin and Strandberg type are preferably used.

[0144] The tungsten precursors that 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 possibly silicotungstic acid (H4SiWi2O40) and its salts. The sources of tungsten can also be any heteropolycompound of the Keggin, vacated Keggin, substituted Keggin, Dawson type, for example. Preferably, ammonium oxides and salts such as ammonium metatungstate or heteropolyanions of the Keggin, vacated Keggin or substituted Keggin type are used.

[0145] 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.

[0146] The nickel precursors which can be used are advantageously chosen from oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example.

[0147] The preferred phosphorus precursor is orthophosphoric acid H3PO4, but its salts and esters such as ammonium phosphates are also suitable. Phosphorus may also be introduced together with the group VIB element(s) in the form of Keggin, vacated Keggin, substituted Keggin or Strandberg type heteropolyanions.

[0148] 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.Preferably, the drying step has a duration of between 5 minutes and 15 hours, preferably between 30 minutes and 12 hours.

[0149] 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 above 200°C under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.

[0150] According to another preferred variant of the invention, 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°C and 1000°C and preferably between 200°C and 750°C, for a duration typically between 15 minutes and 10 hours, under an inert atmosphere or under an atmosphere containing oxygen, in the presence of water or not.

[0151] Before contacting 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 sulforeducing medium, that is to say in the presence of H2S and hydrogen, in order to transform the metal oxides into sulfides such as, for example, MoS2, Co9S8 or Ni3S2. The sulfurization is carried out by injecting onto the catalyst a stream containing H2S and hydrogen, or a sulfur compound capable of decomposing into H2S in the presence of the catalyst and hydrogen. Polysulfides such as dimethyl disulfide (DMDS) are H2S precursors commonly used to sulfurize catalysts. The sulfur can also come from the feedstock. The temperature is adjusted so that the H2S reacts with the metal oxides to form metal sulfides.This sulfurization can be carried out in situ or ex situ (inside or outside the reactor) of the reactor of the process according to the invention at temperatures between 200°C and 600°C, and more preferably between 300°C and 500°C.

[0152] 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:

[0153] (S / metal)catalyst > 0.6 x (S / metal)theoretical

[0154] in which:

[0155] (S / metal)catalyst is the molar ratio between sulfur (S) and metal present on the catalyst

[0156] (S / metal) theoretical is the molar ratio between sulfur and metal corresponding to the total sulfurization of the metal into sulfide.

[0157] This theoretical molar ratio varies according to the metal considered:

[0158] - (S / Fe)theoretical = 1 (S / Co)theoretical— 1 (S / Ni)theoretical — 1 - (S / Mo)theoretical=2 / 1 - (S / W)theoretical=2 / 1

[0159] 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 sulfurization of each metal into sulfide, the calculation being carried out in proportion to the relative molar fractions of each metal.

[0160] Schemes that can be implemented within the framework of the invention

[0161] Different schemes can be implemented in order to produce, at lower cost, a desulfurized and mercaptan-reduced gasoline. The choice of the optimal scheme depends on the characteristics of the gasolines to be processed and produced as well as the constraints specific to each refinery.

[0162] The diagrams described below are given for illustration purposes in a non-limiting manner.

[0163] With reference to [Fig. 1] and according to an embodiment of the method 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 (step a)) in order to selectively hydrogenate the diolefins and to make the light mercaptans heavier. The effluent with a low diolefin and mercaptan content is withdrawn from the reactor 2 via line 4 and is sent to a fractionation column 5 (step b)) configured to separate the effluent and obtain three distinct cuts: a gaseous fraction 8 which is mainly made up of hydrogen in excess and possibly light hydrocarbons Cl to C4, a light gasoline cut 10 (or light gasoline) and a heavy gasoline cut 11 which is made up of the complementary heavy fraction of light gasoline.

[0164] A condenser (not shown) then a gas / liquid separator 7 makes it possible to condense and separate the C5+ compounds entrained by stripping in the gaseous phase evacuated via line 6 to create a liquid reflux 9 which is reinjected in its entirety at the top of the column. The light gasoline cut 10 is withdrawn in the upper part of the column in order to limit its vapor pressure, at the theoretical stage 27; the column comprising 30 theoretical stages, the theoretical stage 1 being the reboiler and the stage 30 being the condenser. The sulfur content of the light gasoline cut is less than 10 ppm by weight, relative to the total weight of the light gasoline cut. The light gasoline cut obtained can be directly sent to the gasoline pool as fuel.

[0165] The heavy gasoline fraction and hydrogen are then sent via line 11 to the hydrodesulfurization unit 13 (step c)). The hydrodesulfurization unit 13 of step c) is, for example, a reactor containing a supported hydrodesulfurization catalyst based on a metal from group VIII and VIB in a fixed bed or in a fluidized bed; preferably, a fixed bed reactor is used. The reactor is operated under operating conditions and in the presence of a hydrodesulfurization catalyst, as described above to decompose the sulfur compounds and form hydrogen sulfide (H2S). During the hydrodesulfurization in step c), recombinant mercaptans are formed by the addition of H2S formed on the olefins. The effluent from the hydrodesulfurization unit 13 is then introduced into the so-called finishing hydrodesulfurization unit 15 (step d) via line 14 without removing the H2S formed.The hydrodesulfurization unit 15 of step d) 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. The unit 15 is operated at a higher temperature than the unit 13 and in the presence of a selective catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII to decompose at least in part the recombinant mercaptans into olefins and H2S. It also makes it possible to hydrodesulfurize the more refractory sulfur compounds.

[0166] The desulfurized heavy gasoline cut is sent via line 16 to a separation tank 17 (step e)) in order to withdraw a gas phase containing H2S and hydrogen via line 18 and a liquid fraction via line 19. The liquid fraction which contains the desulfurized heavy gasoline as well as a fraction of the dissolved H2S is sent via line 19, possibly mixed with the light gasoline cut coming from line 10, to a stabilization column or debutanizer 20 in order to separate at the top of the column via line 21 a stream containing C4 hydrocarbons and the residual H2S and at the bottom of the column via line 22, i.e. the heavy gasoline. desulfurized and stabilized, i.e. the mixture of desulfurized and stabilized light and heavy cut species.

[0167] The following examples illustrate the invention without limiting its scope. Examples

[0168] The analysis methods used to characterize the loads and effluents are as follows:

[0169] - density according to the NF EN ISO 12185 method;

[0170] - sulfur content according to ASTM D2622 method for contents above 10 ppm S and ISO 20846 for contents below 10 ppm S;

[0171] - mercaptan content according to ASTM D3227 method;

[0172] - distillation according to the CSD method simulated distillation "CSD according to the method ASTM2887;

[0173] - diolefin content, determined via the anhydride index maleic (MAV or “Maleic Anhydride Value” according to Anglo-Saxon terminology), according to the UOP 326 method; Example 1 (not in accordance with the invention)

[0174] Table 1 gives the characteristics of an FCC gasoline treated by the process according to [Fig.2]. The FCC gasoline (line 1) is treated in the selective hydrogenation reactor 2 in the presence of a catalyst A. Catalyst A is a NiMo on alumina type catalyst. The metal contents are respectively 7% by weight NiO and 11% by weight MoO3 relative to the total weight of the catalyst, i.e. a Ni / Mo molar ratio of 1.2. The specific surface area of ​​the catalyst is 230 m2 / g. Prior to its use, catalyst A is sulfurized at atmospheric pressure in a sulfurization bench under a H2S / H2 mixture consisting of 15% by volume of H2S at 1 L / gh of catalyst and at 400°C for two hours. This protocol makes it possible to obtain a sulfurization rate greater than 80%.

[0175] Gasoline (line 1) is brought into contact with hydrogen (line 3) in a reactor containing catalyst A. This step of the process carries out the selective hydrogenation of diolefins and the conversion (heaviness) of a portion of the light mercaptan compounds (RSH) to C5- present in the feedstock. The diolefin content is directly proportional to the MAV value (maleic anhydride index or Maleic Anhydrid Value according to English terminology). Diolefins are undesirable compounds because they are precursors of gums in gasoline.

[0176] The operating conditions implemented in the selective hydrogenation reactor are: Temperature: 150°C, Total pressure: 2.5 MPa, volume ratio of added H2 / gasoline charge: 5 normal m3 of hydrogen per m3 of gasoline at standard conditions (vol / vol), hourly volumetric flow rate (WH): 3 h1.

[0177] [Tableauxl] Line 1 Feed Line 4 Selective hydrogenation effluent Organic sulfur content (ppm weight S) 518 512 MAV (mg / g) 13.3 2.1 Olefin content (% weight) 50.4% 49.1% Light mercaptans content C5- (ppm weight) 27 1.55 Simulated distillation (ASTM D2887) 5% distilled mass (°C) 23 23 50% distilled mass (°C) 73 73 95% distilled mass (°C) 145 145

[0178] Table 1: Characteristics of the feedstock (1) and the selective hydrogenation effluent (4).

[0179] The effluent from the selective hydrogenation stage (line 4) with a low content of conjugated diolefins (MAV = 2.1 mg / g) and a low content of light mercaptans (made heavier in the selective hydrogenation stage) is sent to a fractionation column 5 in order to produce a light gasoline (line 10) withdrawn at the reflux drum of the column, a heavy gasoline cut (line 11) at the bottom of the column and a gas stream (line 8) containing essentially hydrogen and light C1-C4 hydrocarbons. The column operates at a pressure of 0.6 MPa and comprises 30 theoretical plates.The gas stream is obtained by condensation at a temperature of 65°C of the gas phase produced at the top of the fractionation column (line 6) then by gas / liquid separation which allows the elimination of light hydrocarbons and excess hydrogen from step a), and the production of a liquid phase (line 9) consisting of C5 hydrocarbons entrained by stripping at the top of the column which is partly returned as reflux into the column.

[0180] The characteristics of the light gasoline cut and the heavy gasoline cut are indicated in Table 2. The light gasoline obtained (line 10) has a low sulfur content (less than 10 ppm by weight). The heavy gasoline cut corresponds to approximately 90% by mass of the gasoline, has a high sulfur content and requires additional treatment before being incorporated into the gasoline pool.

[0181] [Tables2] Line 10 Light Gasoline Line 11 Heavy Gasoline Line 8 Gas Purge Mass Percentage of Cut (%) 7.7 91.6 0.7 Organic Sulfur Content (ppm wt S) 7,558 - Olefin Content (wt%) 56.2% 48.5% - C5+ Loss in Purge (mol%) - - 0.45 Simulated Distillation (ASTM D2887) 5% Distilled Mass (°C) -25.8 11.0 N / A* 30% Distilled Mass (°C) 4.4 52.0 N / A* 50% Distilled Mass (°C) 25.6 71.6 N / A* 70% Distilled Mass (°C) 32.6 100.3 95% Distilled Mass (°C) 43.0 158.3 N / A*

[0182] *N / A = Not Applicable

[0183] Table 2: Characteristics of light and heavy gasoline cuts after the gasoline fractionation stage

[0184] The light gasoline (line 10) produced according to the example not in accordance with the invention vaporizes at atmospheric pressure and at 20°C (70% vaporization weight). An additional column to remove dissolved hydrogen is therefore necessary before to supply a storage of this light gasoline cut. Example 2 (not in accordance with the invention)

[0185] Table 3 gives the characteristics of an FCC gasoline treated by the process according to [Fig.2]. The FCC gasoline (line 1) is treated in the selective hydrogenation reactor 2 in the presence of a catalyst B. Catalyst B is a NiMo on alumina type catalyst. The metal contents are respectively 7% by weight NiO and 11% by weight MoO3 relative to the total weight of the catalyst, i.e. a Ni / Mo molar ratio of 1.2. The specific surface area of ​​the catalyst is 68 m2 / g. Prior to its use, catalyst B is sulfurized at atmospheric pressure in a sulfurization bench under a H2S / H2 mixture consisting of 15% by volume of H2S at 1 L / gh of catalyst and at 400°C for two hours. This protocol makes it possible to obtain a sulfurization rate greater than 80%.

[0186] Gasoline (line 1) is brought into contact with hydrogen (line 3) in a reactor containing catalyst B. This step of the process carries out the selective hydrogenation of diolefins and the conversion (heaviness) of a portion of the light mercaptan compounds (RSH) to C5- present in the feedstock. The diolefin content is directly proportional to the MAV value (maleic anhydride index or Maleic Anhydrid Value according to English terminology). Diolefins are undesirable compounds because they are precursors of gums in gasoline.

[0187] The operating conditions implemented in the selective hydrogenation reactor are: Temperature: 150°C, Total pressure: 2.5 MPa, volume ratio of added H2 / gasoline charge: 5 normal m3 of hydrogen per m3 of gasoline at standard conditions (vol / vol), hourly volumetric flow rate (WH): 3 h1.

[0188] [Tables3] Line 1 Feed Line 4 Selective hydrogenation effluent Organic sulfur content (ppm weight S) 518 512 MAV (mg / g) 13.3 5.1 Olefin content (% weight) 50.4% 49.6% Light mercaptans content C5- (ppm weight) 27 2.42 Simulated distillation (ASTM D2887) 5% distilled mass (°C) 23 23 50% distilled mass (°C) 73 73 95% distilled mass (°C) 145 145

[0189] Table 3: Characteristics of the feedstock (1) and the selective hydrogenation effluent (4).

[0190] The effluent from the selective hydrogenation stage (line 4) with a lower content of conjugated diolefins (MAV = 2.1 mg / g) and a lower content of light mercaptans (made heavier in the selective hydrogenation stage) is sent to a fractionation column (5) in order to produce a light gasoline (line 10) withdrawn 5 trays below the column head, a heavy gasoline cut (line 11) at the bottom of the column and a gas stream (line 8) containing essentially hydrogen and light C1-C4 hydrocarbons. The column operates at a pressure of 0.6 MPa and comprises 30 theoretical trays.The gas stream is obtained by condensation at a temperature of 65°C of the gas phase produced at the top of the fractionation column (line 6) then by gas / liquid separation which allows the elimination of light hydrocarbons and excess hydrogen from step a), and the production of a liquid phase (line 9) consisting of C5 hydrocarbons entrained by stripping at the top of the column which is partly returned as reflux into the column.

[0191] The characteristics of the light gasoline cut and the heavy gasoline cut are indicated in Table 4. The light gasoline obtained (line 10) has a sulfur content greater than 10 ppm by weight and does not allow its incorporation into the gasoline pool. The heavy gasoline cut corresponds to approximately 70% by mass of the gasoline, has a high sulfur content and requires additional treatment before being incorporated into the gasoline pool.

[0192] [Tables4] Line 10 Light Gasoline Line 11 Heavy Gasoline Line 8 Gas Purge Mass Percentage of Cut (%) 29 69.5 1.5 Organic Sulfur Content (ppm wt S) 12,731 - Olefin Content (wt%) 63.3% 43.7% - C5+ Loss in Purge (mol%) - - 0.86 Simulated Distillation (ASTM D2887) 5% Distilled Mass (°C) -19.0 36.4 N / A* 30% Distilled Mass (°C) 22.7 64.7 N / A* 50% Distilled Mass (°C) 32.9 91.8 N / A* 70% Distilled Mass (°C) 40.2 113.7 95% Distilled Mass (°C) 74.3 157.4 N / A*

[0193] *N / A = Not Applicable

[0194] Table 4: Characteristics of light and heavy gasoline cuts after the gasoline fractionation stage Example 3 (in accordance with the invention)

[0195] This example refers to the present invention, according to [Fig. 1]. The selective hydrogenation step is carried out under the same conditions as Example 1, and using the same selective hydrogenation catalyst as Example 1, i.e. NiMo on alumina type catalyst A with a specific catalyst surface area of ​​230 m2 / g.

[0196] The effluent from the selective hydrogenation step of Example 1 (line 4) with a low content of conjugated diolefins (MAV = 2.1 mg / g) and a low content of light sulfur compounds (made heavier in the selective hydrogenation step) is sent to a fractionation column 5 in order to produce a light gasoline (line 10) withdrawn 5 trays below the column head, a heavy gasoline cut (line 11) at the bottom of the column and a gas stream (line 8) containing essentially hydrogen and light C1-C4 hydrocarbons. The column operates at a pressure of 0.6 MPa. The column has 30 theoretical trays. The reboiling power is identical to that of Example 1.The gas stream is obtained by condensation at a temperature of 65°C of the gas phase produced at the top of the fractionation column (line 6) then by gas / liquid separation which allows the elimination of light hydrocarbons and excess hydrogen from step a), and the production of a liquid phase (line 9) consisting of C5 hydrocarbons entrained by stripping at the top of the column which is returned as reflux to the column.

[0197] The characteristics of the light gasoline cut and the heavy gasoline cut are shown in Table 5. The light gasoline obtained (line 10) has a low sulfur content (less than 10 ppm by weight). The heavy gasoline cut, which corresponds to approximately 70% by mass of the gasoline, has a high sulfur content (728 ppm) and requires additional treatment before being incorporated into the gasoline pool.

[0198] [Tables5] Line 10 Light Gasoline Line 11 Heavy Gasoline Line 8 Gas Purge Mass Percentage of Cut (%) 29 70.2 0.8 Organic Sulfur Content (ppm wt S) 9,728 - Olefin Content (wt%) 63.2% 43.1% - C5+ Loss in Purge (mol%) - - 0.44 Simulated Distillation (ASTM D2887) 5% Distilled Mass (°C) -19.0 36.4 N / A* 30% Distilled Mass (°C) 22.9 64.2 N / A* 50% Distilled Mass (°C) 32.0 91.8 N / A* 70% Distilled Mass (°C) 40.3 113.7 95% Distilled Mass (°C) 70.4 157 N / A*

[0199] *N / A = Not Applicable

[0200] Table 5: Characteristics of the cuts: Light gasoline and heavy gasoline after step b) according to the invention

[0201] The light gasoline (line 10) produced according to example 3 does not vaporize at atmospheric pressure and at 20°C and can therefore directly supply storage.

[0202] The process according to Example 3 makes it possible to obtain a light gasoline cut with a low sulfur content while maximizing the olefin content in this cut. The quantity of olefins in the heavy gasoline cut is therefore minimized in order to limit octane loss.

Claims

Claims

1. Process for producing a light gasoline comprising a sulfur content of less than 10 ppm by weight relative to the total weight of said light gasoline, from a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps: a) a selective hydrogenation step so as to hydrogenate the diolefins and carry out a reaction to make a portion of the sulfur compounds heavier, in which process the gasoline and hydrogen are brought into contact with a selective hydrogenation catalyst, at a temperature of between 100°C and 220°C, with a liquid space velocity of between 1 h 1 and 7 h 1 and a pressure of between 0.5 MPa and 5 MPa, and with a molar ratio between hydrogen and the diolefins to be hydrogenated of greater than 1 mol / mol and less than 100 mol / mol,with a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of feed to be treated expressed in m3 per hour under standard conditions of between 2 Nm3 / m3 and 100 Nm3 / m3, said selective hydrogenation catalyst comprising an active phase containing at least one metal from group VIB and at least one metal from group VIII, and a porous support containing at least alumina, said catalyst comprising a specific surface area of ​​between 100 m2 / g and 400 m2 / g; b) a step of separating the effluent obtained at the end of step a) into a gaseous fraction, a light gasoline cut and a heavy gasoline cut, said step b) being carried out in a fractionation column comprising n trays, n being an integer greater than or equal to 20, the first tray being the reboiler and tray “n” being the condenser, it being understood that the light gasoline cut is withdrawn from said fractionation column at tray “ni” with i between 1 and 10.,

2. The method of claim 1, wherein said fractionation column comprises between 20 and 100 trays.

3. A method according to any one of claims 1 or 2, wherein said light gasoline cut is withdrawn from said fractionation column at the “ni” tray with i between 1 and 6.

4. A method according to any one of claims 1 to 3, wherein said Group VIII metal is selected from nickel.

5. A method according to any one of claims 1 to 4, wherein said Group VIB metal is molybdenum.

6. Process according to any one of claims 1 to 5, in which step a) is carried out in the presence of a catalyst comprising nickel with a content by weight of nickel oxide, in NiO form, of between 1% and 12%, and molybdenum with 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.

7. A method according to any one of claims 1 to 6, wherein said catalyst comprises a specific surface area of ​​between 100 m2 / g and 280 m2 / g.

8. Process according to any one of claims 1 to 7, which process further comprising a step c) of hydrodesulfurization of the heavy gasoline HCN cut obtained at the end of step b) in the presence of hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase comprising a metal from group VIB and a metal from group VIII, at a temperature between 210°C and 320°C, at a pressure between 1 MPa and 4 MPa, with a space velocity between 1 h 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of feed to be treated expressed in m3 per hour at standard conditions between 100 Nm3 / m3 and 600 Nm3 / m3, so as to convert at least a portion of the sulfur compounds into H2S.

9. The method of claim 8, wherein the hydrodesulfurization catalyst of step c) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said hydrodesulfurization catalyst containing a content by weight relative to the total catalyst weight of cobalt oxide, in CoO form, of between 0.1 and 10%, a content by weight relative to the total catalyst weight of molybdenum oxide, in MoO3 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 catalyst weight of phosphorus oxide in P2O5 form of between 0.3 and 10% when phosphorus is present, said catalyst having a specific surface area of ​​between 50 m2 / g and 250 m2 / g.

10. A process according to any one of claims 8 or 9, which process further comprises a step d) of hydrodesulfurization of finishing the effluent obtained from step c) without elimination of the H2S formed, in the presence of hydrogen and a hydrodesulfurization catalyst comprising an oxide support and an active phase consisting of at least one metal from group VIII, at a temperature between 280°C and 400°C, at a pressure between 0.5 MPa and 5 MPa, with a space velocity between 1 h 1 and 10 h 1 and a ratio between the hydrogen flow rate expressed in normal m3 per hour and the flow rate of feed to be treated expressed in m3 per hour under standard conditions between 100 Nm3 / m3 and 600 Nm3 / m3;

11. A method according to claim 10, wherein the hydrodesulfurization catalyst of step d) consists of alumina and nickel, said hydrodesulfurization catalyst containing a content by weight relative to the total weight of catalyst of nickel oxide, in NiO form, of between 5 and 20%, said hydrodesulfurization catalyst having a specific surface area of ​​between 30 m2 / g and 180 m2 / g.

12. Method according to any one of claims 10 or 11, further comprising a step e) of separating the H2S formed and present in the effluent from step d).

13. A process according to any one of claims 1 to 12, wherein the gasoline is a catalytic cracked gasoline.

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