CN118434832A

CN118434832A - Method for treating gasoline containing sulfur compounds comprising a dilution step

Info

Publication number: CN118434832A
Application number: CN202280084779.2A
Authority: CN
Inventors: S·库德克; M·德林格; A·戈梅; M-C·马里昂; A·费康特
Original assignee: IFP Energies Nouvelles IFPEN
Current assignee: IFP Energies Nouvelles IFPEN
Priority date: 2021-12-20
Filing date: 2022-12-12
Publication date: 2024-08-02
Also published as: AR128007A1; WO2023117533A1; FR3130835A1

Abstract

A method of treating gasoline containing sulfur compounds, olefins and diolefins, said method comprising the steps of: a) Contacting gasoline, hydrogen and a hydrodesulfurization catalyst to obtain a partially desulfurized effluent; b) Directly contacting the partially desulphurised effluent obtained at the end of step a) with a gaseous or liquid feedstock as diluent under standard temperature and pressure conditions without any separation, to obtain a partially desulphurised and diluted effluent; c) Contacting the partially desulphurised and diluted effluent obtained at the end of step b), hydrogen and catalyst to obtain a desulphurised effluent.

Description

Method for treating gasoline containing sulfur compounds comprising a dilution step

Technical Field

The present invention relates to a process for producing gasoline having low sulfur and low mercaptan content.

Prior Art

The production of gasoline meeting new environmental standards requires a significant reduction in its sulfur content.

It is also known that converted gasolines, more particularly those from catalytic cracking, can account for 30% to 50% of the gasoline pool, with high levels of mono-olefins and sulfur.

Thus, nearly 90% of the sulfur present in gasoline is gasoline resulting from catalytic cracking processes, which is hereinafter referred to as FCC (fluid catalytic cracking) gasoline. Thus, FCC gasoline is a preferred feedstock for the process of the present invention.

Among the possible routes for producing low sulfur content fuels, a widely adopted route includes the specific treatment of sulfur-rich gasoline base stock by catalytic hydrodesulfurization in the presence of hydrogen. Conventional processes desulfurize gasoline in a non-selective manner by hydrogenating a substantial portion of the mono-olefins, which results in large octane number losses and large hydrogen consumption. More recent processes, such as the Prime G+ (trademark) process, can desulfurize cracked gasoline enriched in olefins while limiting the hydrogenation of mono-olefins, thereby limiting the resulting octane number loss and high hydrogen consumption. Such methods are described, for example, in patent applications EP1077247 and EP 1174485.

Patent application US2009065396 discloses a multi-step hydrodesulphurisation process such that the second step is operated at a higher temperature.

Documents EP1857527 and WO2001/38457 disclose two-step hydrodesulphurisation processes with an intermediate step of extracting the H ₂ S formed during the first step.

Documents US5985136 and WO2003/099963 disclose two-step processes in which the catalysts used in the first and second steps are different.

However, there remains a need to limit the loss of olefin hydrogenation during the hydrodesulfurization of gasoline feedstock to achieve sulfur specifications.

Subject of the invention

The object of the present invention is to implement a process for producing low sulfur gasoline which allows upgrading of all sulfur-containing gasoline fractions, preferably catalytically cracked gasoline fractions, and reducing the sulfur content of said gasoline fractions to very low levels without substantially reducing the gasoline yield, while minimizing the octane number reduction caused by olefin hydrogenation.

The subject of the present invention is therefore a process for treating gasolines containing sulfur compounds, olefins and diolefins, comprising at least the following steps:

a) Contacting gasoline, hydrogen and a hydrodesulphurisation catalyst comprising an oxide support and an active phase comprising a group VIB metal and a group VIII metal at a temperature of 210 to 320 ℃, at a pressure of 1 to 4MPa, in a ratio of a space velocity of 1 to 10h ^-1 and a hydrogen flow rate of 100 to 600Nm ³/m³ (expressed as standard m ³/h) to the flow rate of the feedstock to be treated (expressed as standard m ³/h) to obtain a partially desulphurised effluent;

b) The partially desulphurised effluent obtained at the end of step a) is directly contacted with a gaseous or liquid feedstock as diluent without any separation, under standard temperature and pressure conditions (defined herein as temperature 15 ℃ (288.15K) and pressure 0.1 MPa), to obtain a partially desulphurised and diluted effluent;

c) Contacting the partially desulphurised and diluted effluent obtained at the end of step b) with a hydrodesulphurisation catalyst comprising an oxide support and an active phase comprising a group VIB metal and a group VIII metal, at a temperature of 210 to 320 ℃, at a pressure of 1 to 4MPa, at a space velocity of 1 to 10h ^-1, to obtain a desulphurised effluent.

The applicant has unexpectedly found that the intermediate step of diluting the gasoline feedstock to be treated, carried out between two hydrodesulphurisation steps, can significantly improve the stability of the hydrodesulphurisation performance by controlling the exothermicity of the olefin hydrogenation, so that the deactivation phenomenon of the hydrodesulphurisation catalyst can be limited.

According to one or more embodiments, the diluent of step b) is a gaseous feedstock comprising at least 50% by volume of hydrogen relative to the total volume of the gaseous feedstock.

According to one or more embodiments, the ratio of the volumetric flow rate of the gaseous feed introduced in step b) to the volumetric flow rate of the hydrogen introduced in step a) is from 0.01 to 25Nm ³/Nm³.

According to one or more embodiments, the diluent of step b) is a liquid hydrocarbon-based feedstock having a boiling point of from 30 ℃ to 250 ℃.

According to one or more embodiments, the ratio of the volumetric flow rate of the liquid feedstock introduced in step b) to the volumetric flow rate of the gasoline feedstock introduced in step a) is from 0.01 to 1Sm ³/Sm³ (standard cubic meter/standard cubic meter).

According to one or more embodiments, the liquid feedstock introduced in step b) is the same as the gasoline feedstock introduced in step a) of the process.

According to one or more embodiments, the gaseous feed or the liquid feed is introduced in step b) at a temperature lower than the temperature of the partially desulphurised effluent obtained at the end of step a).

According to one or more embodiments, the gaseous feed or the liquid feed is introduced in step b) at a temperature of 20 ℃ to 300 ℃.

According to one or more embodiments, steps a) and b) are performed in the same reactor.

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

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

According to one or more embodiments, the catalyst of step a) and/or step c) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, the catalyst comprising cobalt oxide in the form of CoO in an amount of from 0.1% to 10% relative to the total weight of the catalyst, molybdenum oxide in the form of MoO ₃ in an amount of from 1% to 20% relative to the total weight of the catalyst, the cobalt/molybdenum molar ratio being from 0.1 to 0.8, phosphorus oxide in the form of P ₂O₅ in an amount of from 0.3% to 10% relative to the total weight of the catalyst when phosphorus is present, the catalyst having a specific surface area of from 50 to 250m ²/g.

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

According to one or more embodiments, the gasoline is contacted with hydrogen and a selective hydrogenation catalyst prior to step a) and prior to any optional distillation step, to selectively hydrogenate the diolefins contained in the gasoline to olefins.

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

List of drawings

Fig. 1 is a schematic diagram of an embodiment according to the present invention.

Detailed Description

Definition of the definition

Hereinafter, the groups of chemical elements are given according to the CAS taxonomy (CRC Handbook of CHEMISTRY AND PHYSICS, CRC Press, main edition D.R.Limde, 81 th edition, 2000-2001). For example, group VIII according to CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification.

The term "specific surface area" is understood to mean The BET specific surface area (S _BET, in m ²/g) determined by nitrogen adsorption according to The standard ASTM D3663-78 established by The Brunauer-Emmett-Teller method described in Journal of THE AMERICAN CHEMICAL Society, 1938, 60, 309.

The total pore volume of the catalyst or of the support used for preparing the catalyst is understood to mean that a surface tension of 484 dynes/cm and a contact angle of 140℃are used at a maximum pressure of 4000 bar (400 MPa) by mercury intrusion according to the standard ASTM D4284, for exampleThe instrument, model number Autopore III, measured volume.

According to the proposal of pages 1050-1055 of publication "Techniques del′ingénieur,traité analyse et caractérisation"[Techniques of the Engineer,Analysis and Characterization Treatise] written by Jean Charpin and Bernard Rasneur, a wetting angle equal to 140 ° is used. For better accuracy, the value of the total pore volume corresponds to the value of the total pore volume measured by mercury porosimetry for a sample minus the value of the total pore volume measured by mercury porosimetry for the same sample at a pressure of 30psi (about 0.2 MPa).

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

Description of the raw materials

The process according to the invention can treat any type of gasoline fraction containing sulphur compounds and olefins, for example fractions produced by coking, visbreaking, steam cracking or catalytic cracking (FCC, fluid catalytic cracking) units. The gasoline may contain large amounts of gasoline from other production processes, such as atmospheric distillation (gasoline from direct distillation (or straight run gasoline)), or from conversion processes (coker or steam cracker gasoline). The feedstock preferably consists of a gasoline fraction from a catalytic cracking unit.

The feedstock is a gasoline fraction containing sulfur compounds and olefins, typically having a boiling point in the range of from the boiling point of a hydrocarbon having 2 or 3 carbon atoms (C2 or C3) to 260 ℃, preferably from the boiling point of a hydrocarbon having 2 or 3 carbon atoms (C2 or C3) to 220 ℃, more preferably from the boiling point of a hydrocarbon having 5 carbon atoms to 220 ℃. The process according to the invention also makes it possible to treat feedstocks having a lower final boiling point than described above, for example C5-180℃ fractions.

The sulfur content of a gasoline fraction produced by catalytic cracking (FCC) depends on the sulfur content of the FCC treated feedstock, whether the FCC feedstock has been pretreated, and the final boiling point of the fraction. Typically, the sulfur content of the whole gasoline fraction (especially the gasoline fraction derived from FCC) is greater than 100 ppm by weight, most often greater than 500 ppm by weight. For gasolines with a final boiling point of greater than 200 ℃, the sulfur content is generally greater than 1000 ppm by weight; in some cases they may even be on the order of 4000 to 5000 ppm by weight.

The feedstock treated by the process of the present invention may be a feedstock containing sulfur compounds having a sulfur content of greater than 1000 ppm by weight and typically greater than 1500 ppm.

In addition, the gasoline produced by a catalytic cracking (FCC) unit contains on average 0.5 to 5 wt% diolefins, 20 to 50wt% olefins and 10 to 0.5 wt% sulfur, typically containing less than 300ppm mercaptans.

Selective hydrogenation step a 0) (optional)

Depending on the type of gasoline to be treated, it may be advantageous to treat the gasoline in advance in the presence of hydrogen and a selective hydrogenation catalyst to at least partially hydrogenate the diolefins and to carry out the reaction of increasing the molecular weight of the fraction of light mercaptan (RSH) compounds present in the feedstock by reaction with the olefins to form thioethers.

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 increasing the molecular weight of the light mercaptans. The reaction of selectively hydrogenating the diolefins and increasing the molecular weight of the light mercaptans is preferably carried out over a sulfided catalyst comprising at least one group VIII element and optionally at least one group VIB element and an oxide support. The group VIII element is preferably selected from nickel and cobalt, in particular nickel. The group VIB element (when present) is preferably selected from molybdenum and tungsten, very preferably molybdenum.

The oxide support of the catalyst is preferably selected from alumina, nickel aluminate, silica, silicon carbide or mixtures of these oxides. Preferably, alumina is used, more preferably, high purity alumina is used. According to a preferred embodiment, the selective hydrogenation catalyst comprises nickel in an amount of from 1% to 12% by weight (calculated as NiO) of nickel oxide and molybdenum in an amount of from 6% to 18% by weight (calculated as MoO ₃) of molybdenum oxide, with a nickel/molybdenum molar ratio of from 0.3 to 2.5, said metal being deposited on a support consisting of alumina. The sulfiding degree of the metal constituting the catalyst is preferably more than 60%.

In the optional selective hydrogenation step, the gasoline is contacted with the catalyst at a temperature of from 50 to 250 ℃, preferably from 80 to 220 ℃, more preferably from 90 to 200 ℃, a liquid space velocity (LHSV) in units of feed volume per volume of catalyst bed per hour (l/l/h) of from 0.5h ^-1 to 20h ^-1. The pressure is 0.4 to 5MPa, preferably 0.6 to 4MPa, more preferably 1 to 3MPa. The optional selective hydrogenation step is generally carried out in a ratio of a hydrogen flow (expressed in m ³ per hour) of from 2 to 100Nm ³/m³, preferably from 3 to 30Nm ³/m³, to a feed flow (expressed in m ³ per hour under standard conditions) to be treated.

After the selective hydrogenation, the diene content, as determined by the Maleic Anhydride Value (MAV) according to the UOP 326 method, is generally reduced to less than 6mg maleic anhydride/g, even less than 4mg MA/g, more preferably less than 2mg MA/g. In some cases, it is possible to obtain a diene content of less than 1mg MA/g.

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

It should be noted that it is conceivable to simultaneously carry out the step of hydrogenating the diolefins and the step of fractionating into two to three fractions by means of a catalytic distillation column comprising a distillation column equipped with at least one catalytic bed.

Hydrodesulfurization step a)

Hydrodesulfurization step a) is carried out to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H ₂ S.

The temperature is generally from 210 to 320℃and preferably from 220 to 300 ℃. The temperature used must be sufficient to keep the petrol to be treated in the gaseous phase in the reactor.

The operating pressure of this step is generally from 1 to 4MPa, preferably from 1.5 to 3MPa.

The amount of catalyst used in each reactor is generally such that the ratio of the flow rate of the gasoline to be treated, expressed in m ³/h at standard conditions, to the m ³ catalytic bed (also called space velocity or LHSV) is from 1 to 100h ^-1, preferably from 1 to 50h ^-1, very preferably from 2 to 20h ^-1.

The hydrogen flow is generally such that the ratio of the hydrogen flow (expressed as standard m ³/hour (Nm ³/h)) to the feed flow to be treated (expressed as m ³/h at standard conditions (15 ℃ C., 0.1 MPa)) is from 10 to 1000Nm ³/m³, preferably from 50 to 600Nm ³/m³. Standard m ³ is understood to be the volume of 1m ³ gas at 0 ℃ and 0.1 MPa.

The hydrogen required for this step may be fresh hydrogen or recycled hydrogen (which preferably does not contain H ₂ S), or a mixture of fresh hydrogen and recycled hydrogen. Preferably, a mixture of fresh hydrogen and recycled hydrogen is used.

The extent of desulfurization in step a) depends on the sulfur content of the feedstock to be treated, generally being greater than 50%, preferably greater than 70%, so that the product obtained in step a) contains less than 100 ppm by weight of sulfur, preferably less than 50ppm by weight of sulfur.

The catalyst used in step a) must exhibit good selectivity for the hydrodesulphurisation reaction relative to the olefin hydrogenation reaction.

The hydrodesulphurisation catalyst of step a) 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.

The group VIB metal present in the active phase of the catalyst is preferably selected from molybdenum and tungsten. The group VIII metal present in the active phase of the catalyst is preferably selected from cobalt, nickel and mixtures of these two elements. The active phase of the catalyst is preferably selected from the group consisting of nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum, and very preferably the active phase consists of cobalt and molybdenum.

The content of the group VIII metal is 0.1 to 10 wt%, preferably 0.6 to 8wt%, preferably 0.6 to 7wt%, very preferably 1 to 6wt% of the oxide of the group VIII metal relative to the total weight of the catalyst.

The content of group VIB metals is from 1 wt.% to 20 wt.%, preferably from 2 wt.% to 18 wt.%, very preferably from 3 wt.% to 16 wt.% of oxides of group VIB metals, relative to the total weight of the catalyst.

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

Optionally, the catalyst may also have a phosphorus content of generally 0.3 to 10 wt%, preferably 0.3 to 5 wt%, very preferably 0.5 to 3 wt% of P ₂O₅, relative to the total weight of the catalyst. For example, the phosphorus present in the catalyst is combined with a group VIB metal, and optionally also with a group VIII metal, in the form of a heteropolyanion.

Further, when phosphorus is present, the phosphorus/(group VIB metal) molar ratio is usually 0.1 to 0.7, preferably 0.2 to 0.6.

Preferably, the catalyst is characterized by a specific surface area of 5 to 400m ²/g, preferably 10 to 250m ²/g, preferably 50 to 250m ²/g. In the present invention, the specific surface area is determined by the BET method according to standard ASTM D3663, such as Rouquerol f, rouquerol j and Singh k, works Adsorption by Powders & Porous Solids: as described in principles, methodology and Applications, ACADEMIC PRESS,1999, for example by Micromeritics ^TM brand Autopore model III ^TM.

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

The catalyst has a compact bulk density (TPD) of usually 0.4 to 0.8g/ml, preferably 0.4 to 0.7g/ml. TPD measurement involves introducing the catalyst into a measuring cylinder whose volume has been predetermined and then vibrating it by vibration until a constant volume is obtained. The bulk density of the tapped product is calculated by comparing the mass introduced and the volume occupied after tapping.

The catalyst may be in the form of small diameter cylindrical or multilobal (trilobal, tetralobal, etc.) extrudates or spheres.

The oxide support of the catalyst is generally a porous solid selected from: alumina, silica-alumina, titanium or magnesium oxide used alone or in combination with alumina or silica-alumina. Preferably selected from the group consisting of silica, transition alumina family and silica-alumina; very preferably, the oxide support consists essentially of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, even at least 90% by weight, of alumina. It preferably consists of alumina only. Preferably, the oxide support of the catalyst is a "high temperature" alumina, that is, it comprises θ -, δ -, κ -, or α -phase alumina alone or as a mixture, and less than 20 wt% γ -, χ -, or η -phase alumina.

In the process according to the invention, the conversion of unsaturated sulphur compounds is advantageously greater than 15%, preferably greater than 50%. Meanwhile, during this step, the degree of hydrogenation of the olefin is preferably less than 50%, more preferably less than 40%, and very preferably less than 35%.

The partially desulphurised effluent obtained at the end of step a) is then sent directly to step b) of the process of the invention without separation.

A step b) of diluting the effluent of step a)

This step consists in diluting the partially desulphurised effluent obtained at the end of step a) with a gaseous or liquid feedstock as diluent (also referred to herein as "diluent feedstock" or "diluent") under standard temperature and pressure conditions (defined herein as temperature 15 ℃ (288.15K) and pressure 0.1 MPa). Preferably, the diluent feed is gaseous under standard temperature and pressure conditions.

In one embodiment according to the invention, when the diluent is a gaseous feed under standard temperature and pressure conditions, the gaseous feed preferably comprises more than 50% hydrogen H ₂, more preferably more than 70% hydrogen, very preferably more than 90% hydrogen by volume. The volume flow of the gaseous feed is defined such that the ratio of the volume flow of said gaseous feed introduced in step b) to the volume flow of hydrogen introduced in step a) is from 0.01 to 25Nm ³/Nm³, preferably from 0.02 to 4Nm ³/Nm³, very preferably from 0.04 to 1Nm ³/Nm³. The gaseous feed is introduced at a pressure generally in the range of from about 1 to about 4MPa, preferably 1.5 to 3 MPa. The gaseous feed is generally introduced at a temperature below the temperature of the effluent produced in step a), preferably at a temperature of from 20 ℃ to 300 ℃, preferably from 30 ℃ to 280 ℃, more preferably from 30 ℃ to 220 ℃, more preferably from 35 ℃ to 180 ℃, more preferably from 35 ℃ to 120 ℃, and more preferably from 35 ℃ to 80 ℃.

In another embodiment according to the present invention, when the diluent is a liquid feedstock under standard temperature and pressure conditions (defined herein as temperature 15 ℃ (288.15K) and pressure 0.1 MPa), the liquid feedstock may be any liquid hydrocarbon-based feedstock having a boiling point of from 30 ℃ to 250 ℃, preferably from 35 ℃ to 240 ℃, more preferably from 40 ℃ to 220 ℃. Preferably, the liquid feedstock is the same as the gasoline feedstock to be treated introduced during step a) or as the effluent produced in step c). The volumetric flow rate of the liquid feedstock is defined such that the ratio of the volumetric flow rate of the liquid feedstock introduced in step b) to the volumetric flow rate of the gasoline feedstock introduced in step a) is from 0.01 to 1m ³/m³, preferably from 0.02 to 1m ³/m³, very preferably from 0.02 to 0.5m ³/m³. The liquid feedstock is introduced at a pressure generally in the range of from about 1 to about 4MPa, preferably 1.5 to 3 MPa. The liquid feed is introduced at a temperature below the temperature of the effluent produced in step a) and preferably at a temperature of from 20 ℃ to 300 ℃, preferably from 30 ℃ to 280 ℃, very preferably from 40 ℃ to 220 ℃, more preferably from 40 ℃ to 180 ℃, more preferably from 40 ℃ to 120 ℃ and more preferably from 40 ℃ to 80 ℃.

Preferably, step b) is performed in the same reactor as step a).

Hydrodesulfurization step c)

This step comprises converting at least part of the sulphur compounds (e.g. thiophene compounds) contained in the effluent produced in step b) into saturated compounds, such as tetrahydrothiophene (or thiacyclopentane) or mercaptans, or at least partially hydrogenolyzing these sulphur compounds to form H ₂ S.

The amount of catalyst used in each reactor is generally such that the ratio of the flow rate of the gasoline to be treated, expressed in m ³/h at standard conditions, to the m ³ catalytic bed (also called space velocity or LHSV) is from 1 to 100h ^-1, preferably from 1 to 50h ^-1, very preferably from 3 to 20h ^-1.

The hydrogen flow is generally such that the ratio of the hydrogen flow (expressed as standard m ³/hour (Nm ³/h)) to the feed flow to be treated (expressed as m ³/h at standard conditions (15 ℃ C., 0.1 MPa)) is 50 to 1000Nm ³/m³, preferably 100 to 600Nm ³/m³.

The catalyst used in step c) must exhibit good selectivity for the hydrodesulphurisation reaction relative to the olefin hydrogenation reaction.

The hydrodesulphurisation 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.

In a preferred variant, the catalyst used in step c) is the same as the catalyst used in step a).

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

Very preferably, steps a), b) and c) are carried out in the same reactor.

Finishing (finishing) hydrodesulfurization step d) (optional)

During the hydrodesulphurisation steps a) and c), most of the sulphur compounds are converted to H ₂ S. The remaining sulphur compounds are essentially refractory sulphur compounds and recombinant mercaptans resulting from the addition of H ₂ S formed in steps a) and c) to mono-olefins present in the feed.

The "finishing" hydrodesulfurization step is primarily intended to reduce the content of recombinant mercaptans. Preferably, step d) is performed at a higher temperature than steps a) and c). In particular, by using a higher temperature in this step than in steps a) and c), the formation of olefins and H ₂ S is promoted by thermodynamic equilibrium. Step d) also makes it possible to continue the hydrodesulfurization of the residual sulfur compounds.

Hydrodesulfurization step d) comprises contacting the effluent from step c) optionally with added hydrogen in one or more hydrodesulfurization reactors comprising one or more catalysts suitable for carrying out hydrodesulfurization.

The hydrodesulphurisation step d) is carried out without significant hydrogenation of the olefins. The degree of olefin hydrogenation of the catalyst of hydrodesulphurisation step d) is typically less than 5%, more typically less than 2%.

The temperature of this step is typically 280 to 400 ℃, more preferably 290 to 380 ℃, very preferably 300 to 360 ℃. The temperature of step d) is generally at least 5 ℃, preferably at least 10 ℃ and very preferably at least 30 ℃ higher than the temperature of steps a) and c).

The operating pressure of this step is generally from 0.5 to 5MPa, preferably from 1 to 3MPa.

The amount of catalyst used in each reactor is generally such that the ratio of the flow rate of gasoline to be treated expressed in m ³/h at standard conditions to the volume of m ³ catalyst (also called space velocity or LHSV) is from 1 to 10h ^-1, preferably from 2 to 8h ^-1.

The hydrogen flow is generally such that the ratio of the hydrogen flow (expressed as standard m ³/hour (Nm ³/h)) to the feed flow to be treated (expressed as m ³/h under standard conditions (15 ℃ C., 0.1 MPa)) is 50 to 600Nm ³/m³, preferably 50 to 500Nm ³/m³.

The extent of desulfurization in step d) depends on the sulfur content of the feedstock to be treated, generally greater than 50%, preferably greater than 70%, so that the product obtained in step d) contains less than 60 ppm by weight of sulfur, preferably less than 40 ppm by weight of sulfur.

The catalyst of step d) differs in nature and/or composition from the catalysts used in steps a) and c). The catalyst of step d) is in particular a very selective hydrodesulphurisation catalyst: it allows hydrodesulfurization to be performed without hydrogenating the olefins, thereby maintaining the octane number.

Catalysts which can be used in step d) of the process of the invention (this list is not exhaustive) are catalysts comprising an oxide support and an active phase consisting of at least one group VIII metal, and preferably selected from nickel, cobalt and iron. These metals may 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 content of the group VIII metal is 1 to 60 wt%, preferably 5 to 30 wt%, very preferably 5 to 20 wt% of the group VIII metal oxide relative to the total weight of the catalyst.

Preferably, the catalyst is characterized by a specific surface area of 5 to 400m ²/g, preferably 10 to 250m ²/g, preferably 20 to 200m ²/g, very preferably 30 to 180m ²/g. In the present invention, the specific surface area is determined by the BET method according to standard ASTM D3663, such as Rouquerol f, rouquerol j and Singh k, works Adsorption by Powders & Porous Solids: as described in principles, methodology and Applications, ACADEMIC PRESS,1999, for example by Micromeritics ^TM brand Autopore model III ^TM.

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

The catalyst has a compact bulk density (TPD) of usually 0.4 to 0.8g/ml, preferably 0.4 to 0.7g/ml.

TPD measurement involves introducing the catalyst into a measuring cylinder whose volume has been predetermined and then vibrating it by vibration until a constant volume is obtained. The bulk density of the tapped product is calculated by comparing the mass introduced and the volume occupied after tapping.

The oxide support of the catalyst is generally a porous solid selected from: alumina, silica-alumina, titanium or magnesium oxide used alone or in combination with alumina or silica-alumina. Preferably selected from the group consisting of silica, transition alumina family and silica-alumina; very preferably, the oxide support consists essentially of alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, even at least 90% by weight, of alumina. It preferably consists of alumina only. Preferably, the oxide support of the catalyst is a "high temperature" alumina, that is, it comprises θ -, δ -, κ -, or α -phase alumina alone or as a mixture, and less than 20% γ -, χ -, or η -phase alumina.

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

The catalyst of the hydrodesulphurisation step d) is characterized by a hydrodesulphurisation catalytic activity which is generally 1% to 90%, preferably 1% to 70%, very preferably 1% to 50% of the catalytic activity of the catalysts of the hydrodesulphurisation steps a) and c).

Description of catalyst preparation and sulfiding

The preparation of the catalysts of steps a), c) and d) is known and generally comprises the following steps: the active phase in the form of an oxide can be obtained by impregnating the metal of groups VIII and VIB and optionally phosphorus and/or organic compounds, if present, on the oxide support, followed by a drying operation and then an optional calcination. Prior to using the catalyst in a hydrodesulfurization process for sulfur-containing olefinic gasoline fractions, the catalyst is typically sulfided as described below to form the active species.

The impregnation step may be performed by slurry impregnation, overdose impregnation, dry impregnation or any other means known to the person skilled in the art. The impregnating solution is selected so as to be able to dissolve the metal precursor in the desired concentration.

For example, among the molybdenum sources, oxides and hydroxides, molybdic acid and salts thereof, in particular ammonium salts, such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H ₃PMo₁₂O₄₀) and salts thereof, and optionally silicomolybdic acid (H ₄SiMo₁₂O₄₀) and salts thereof, may be used. The molybdenum source may also be any heteropoly compound of the type Keggin, abscission Keggin, substitution Keggin, dawson, anderson, or Strandberg, for example. Molybdenum trioxide and Keggin, abscission Keggin, substituted Keggin and Strandberg type heteropoly compounds are preferably used.

Tungsten precursors that can be used are also well known to those skilled in the art. For example, among tungsten sources, oxides and hydroxides, tungstic acid and salts thereof, in particular ammonium salts, such as ammonium tungstate, ammonium metatungstate, phosphotungstic acid and salts thereof, and optionally silicotungstic acid (H ₄SiW₁₂O₄₀) and salts thereof, may be used. The tungsten source may also be, for example, keggin, abscission Keggin, any heteropoly compound that replaces the Keggin or Dawson type. Oxides and ammonium salts, such as ammonium metatungstate, or Keggin, absent Keggin or substituted Keggin type heteropolyanions, are preferably used.

Useful cobalt precursors are advantageously selected from, for example, oxides, hydroxides, hydroxycarbonates, carbonates and nitrates. Cobalt hydroxide and cobalt carbonate are preferably used.

Useful nickel precursors are advantageously selected from, for example, oxides, hydroxides, hydroxycarbonates, carbonates and nitrates.

The preferred phosphorus precursor is H ₃PO₄, but salts and esters thereof, such as ammonium phosphate, are also suitable. Phosphorus may also be introduced simultaneously with one or more group VIB elements in the form of Keggin, delocalized Keggin, substituted Keggin or Strandberg heteropolyanions.

After the impregnation step, the catalyst is generally subjected to a drying step at a temperature of less than 200 ℃, advantageously between 50 ℃ and 180 ℃, preferably between 70 ℃ and 150 ℃, very preferably between 75 ℃ and 130 ℃. The drying step is preferably carried out under an inert atmosphere or an oxygen-containing atmosphere. The drying step may be carried out by any technique known to those skilled in the art. It is advantageously carried out at atmospheric or reduced pressure. Preferably, this step is carried out at atmospheric pressure. It is advantageously performed in a cross-flow bed (TRAVERSED BED) using hot air or any other hot gas. Preferably, when the drying is carried out in a fixed bed, the gas used is air or an inert gas, such as argon or nitrogen. Very preferably, the drying is carried out in a cross-flow bed in the presence of nitrogen and/or air. Preferably, the duration of the drying step is from 5 minutes to 15 hours, preferably from 30 minutes to 12 hours.

According to a variant of the invention, the catalyst has not been calcined during its preparation, that is to say the impregnated catalytic precursor has not been subjected to a heat treatment step at a temperature higher than 200 ℃ in an inert atmosphere or an oxygen-containing atmosphere, in the presence or absence of water.

According to another preferred variant of the invention, the catalyst is subjected to a calcination step during its preparation, that is to say the impregnated catalytic precursor is subjected to a heat treatment step at a temperature of from 250 ℃ to 1000 ℃, preferably from 200 ℃ to 750 ℃, in the presence or absence of water, in an inert atmosphere or an oxygen-containing atmosphere, generally for a period of from 15 minutes to 10 hours.

The catalyst of the process of the present invention is typically subjected to a sulfiding step prior to contact with the feedstock to be treated in the gasoline hydrodesulfurization process. The sulfidation is preferably carried out in a sulfur reduction medium, i.e. in the presence of H ₂ S and hydrogen, in order to convert the metal oxide into sulfide, such as MoS ₂、Co₉S₈ or Ni ₃S₂. Sulfiding is performed by injecting a stream containing H ₂ S and hydrogen or a stream of sulfur compounds capable of decomposing into H ₂ S in the presence of a catalyst and hydrogen onto the catalyst. Polysulfides, such as dimethyl disulfide (DMDS), are H ₂ S precursors commonly used to sulfid catalysts. Sulfur may also be derived from the feedstock. The temperature is adjusted to react H ₂ S with the metal oxide to form a metal sulfide. The vulcanization may be carried out in situ or ex situ (inside or outside the reactor) of the reactor of the process of the invention, at a temperature of from 200 ℃ to 600 ℃, more preferably from 300 ℃ to 500 ℃.

The metal constituting the catalyst has a degree of sulfidation of at least 60%, preferably at least 80%. The sulfur content in the sulfided catalyst was measured by elemental analysis according to ASTM D5373. A metal is considered sulfided when the total sulfiding degree (defined by the molar ratio of sulfur (S) present on the catalyst to the metal) is at least equal to 60% of the theoretical molar ratio corresponding to complete sulfiding of the metal or metals under consideration. The total degree of vulcanization is defined by the following formula:

(S/metal) _Catalyst is not less than 0.6x (S/metal) _{theoretical value}

Wherein:

(S/Metal) _Catalyst is the molar ratio of sulfur (S) to metal present on the catalyst

(S/metal) _{theoretical value} is the molar ratio of sulfur to metal corresponding to complete sulfidation of the metal to form sulfide.

The theoretical molar ratio varies according to the metal considered:

-(S/Fe)_{theoretical value}＝1

-(S/Co)_{theoretical value}＝1

-(S/Ni)_{theoretical value}＝1

-(S/Mo)_{theoretical value}＝2/1

-(S/W)_{theoretical value}＝2/1

When the catalyst comprises a plurality of metals, the molar ratio of S to the combined metal present on the catalyst must also be at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of each metal to give sulfide, this calculation being made in proportion to the relative mole fraction of each metal.

Other features and advantages of the invention will become apparent from reading the following description, given by way of illustration only and not by way of limitation, with reference to fig. 1.

Referring to fig. 1 and according to one embodiment of the process of the invention, the gasoline to be treated is sent via line 1 and hydrogen is sent via line 2 to the hydrodesulphurisation unit 3 of step a). The hydrodesulphurisation unit 3 of step a) is for example a reactor containing a supported hydrodesulphurisation catalyst based on a group VIII metal and a group VIB metal in a fixed bed or a fluidised bed; preferably, a fixed bed reactor is used. The reactor is operated under the operating conditions described above and in the presence of a hydrodesulfurization catalyst to decompose sulfur compounds and form hydrogen sulfide (H ₂ S). During the hydrodesulphurisation of step a), recombinant mercaptans are formed by adding the formed H ₂ S to olefins. The effluent from hydrodesulfurization unit 3 is then mixed with diluent via line 5 in step b) and introduced into hydrodesulfurization unit 6 via line 4 without removing the H ₂ S formed.

The hydrodesulphurisation unit 6 of step c) is for example a reactor containing a supported hydrodesulphurisation catalyst based on a group VIII metal and a group VIB metal in a fixed bed or a fluidised bed; preferably, a fixed bed reactor is used. The reactor is operated under the operating conditions described above and in the presence of a hydrodesulfurization catalyst to decompose sulfur compounds and form hydrogen sulfide (H ₂ S). During the hydrodesulphurisation of step c), recombinant mercaptans are formed by adding the formed H ₂ S to olefins. The effluent from hydrodesulfurization unit 6 is then introduced via line 7 into a "finishing" hydrodesulfurization unit 8 without removing the H ₂ S formed. The hydrodesulfurization unit 8 of step d) is, for example, a reactor containing a hydrodesulfurization catalyst in a fixed or fluidized bed; preferably, a fixed bed reactor is used. Unit 8 operates at a higher temperature than unit 6 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 the recombinant mercaptans to produce olefins and H ₂ S. It also allows hydrodesulfurization of sulfur compounds that are more difficult to crack.

An effluent, i.e. gasoline containing H ₂ S, compounds boiling above the boiling point of butane, and sulphur compounds including mercaptans, is withdrawn from the hydrodesulphurisation reactor 8 via line 9.

The following examples illustrate the invention without limiting its scope.

Examples

The analytical methods used to characterize the feedstock and effluent were as follows:

-determining density according to NF EN ISO 12185 method;

-for S in an amount greater than 10ppm, determining the sulfur content according to ASTM D2622 method; for S with the content less than 10ppm, the sulfur content is measured according to the ISO 20846 method;

-distillation according to ASTM2887 method, according to CSD simulated distillation method;

-determining the diene content by Maleic Anhydride Value (MAV) according to UOP 326 method.

Example 1 (not according to the invention)

Gasoline (containing 25 weight percent of olefin and 600ppmS total sulfur) obtained from a catalytic cracking unit is subjected to a single hydrodesulfurization step in the presence of a catalyst A in a single catalytic bed of an adiabatic reactor, wherein the catalyst A is an alumina-supported CoMo catalyst, and the metal content is 3 weight percent CoO and 10 weight percent MoO ₃ respectively; the specific surface area of the catalyst was 135m ²/g. The catalyst was sulfided prior to use by contacting it with a feed containing 2 wt% sulfur (as dimethyl disulfide in n-heptane) at 350 ℃ for 4 hours at a pressure of 3.4 MPa.

60Ml of catalyst were loaded into the individual catalytic beds of the reactor.

The operating conditions for the one-step hydrodesulfurization step of a gasoline feedstock are as follows: lhsv=3h ^-1 (relative to the whole catalytic bed), H ₂/HC＝250Nm³/m³, p=2.0 MPa.

Table 1 lists the characteristics of the liquid effluent obtained at the end of the hydrodesulphurisation step and the temperature conditions of the feed stream obtained therefrom over time.

TABLE 1

Time of feed stream	Tiantian (Chinese character of 'Tian')	0	15	30	60	90
							Inlet sulfur content	ppmS	598	598	598	598	598
Outlet sulfur content	ppmS	10	11	11	13	15
							Inlet temperature	℃	266	266	266	266	266
Outlet temperature	℃	288	286	285	284	283

* Olefin hydrogenation

Example 2 (according to the invention)

The gasoline obtained from the catalytic cracking unit, the characteristics of which are described in example 1, is hydrodesulphurised on the sulphurised catalyst a described in example 1. 60ml of catalyst was charged into two catalytic beds (30 ml each) of the reactor.

The overall operating conditions for the hydrodesulfurization of the two catalytic beds are as follows: lhsv=3h ^-1、H₂/HC＝250Nm³/m³, p=2.0 MPa.

In the first catalytic bed, the operating conditions of the first hydrodesulphurisation step are as follows: lhsv=6h ^-1、H₂/HC＝225Nm³/m³.

At the end of this step, a gaseous diluent consisting of hydrogen was added, the ratio of diluent flow/H ₂ flow at the inlet of the first reactor being 0.22Nm ³/Nm³. The diluent is added at a temperature of 50 ℃.

The mixture is then subjected to a second hydrodesulphurisation step on catalyst a in a second catalyst bed. The operating conditions for this hydrodesulfurization step in the second bed are as follows: lhsv=6h ^-1、H₂/HC＝275Nm³/m³.

Table 2 lists the characteristics of the liquid effluent produced in the second hydrodesulfurization step and the temperature conditions of the feed stream obtained therefrom over time.

TABLE 2

Time of feed stream	Tiantian (Chinese character of 'Tian')	0	15	30	60	90
							Inlet sulfur content	ppmS	598	598	598	598	598
Outlet sulfur content	ppmS	10	10	11	12	13
							Bed 1 inlet temperature	℃	269	269	269	269	269
Bed 1 outlet temperature	℃	284	284	283	282	281
							Bed 2 inlet temperature	℃	275	275	274	273	272
Bed 2 outlet temperature	℃	284	284	283	281	279

* Olefin hydrogenation

The exothermic property in this example was lower and the highest temperature was also lower than in example 1. This results in slower deactivation of the catalyst and reduced sulfur content at the outlet of the process of the invention.

Example 3 (according to the invention)

The gasoline obtained from the catalytic cracking unit, the characteristics of which are described in example 1, is hydrodesulphurised on the sulphurised catalyst a described in example 1. 60ml of catalyst were charged into two 30ml beds of the reactor.

The overall operating conditions for the hydrodesulfurization of the total catalyst bed are as follows: lhsv=3h ^-1、H₂/HC＝250Nm³/m³, p=2.0 MPa.

In the first bed, the operating conditions of the hydrodesulfurization step are as follows: lhsv=6h ^-1、H₂/HC＝237.5Nm³/m³.

At the end of this step, a gaseous diluent consisting of hydrogen was added, the ratio of diluent flow/H ₂ flow at the inlet of the first reactor being 0.105Nm ³/Nm³. The diluent is added at a temperature of 50 ℃.

The mixture is then subjected to a second hydrodesulphurisation step on catalyst a in a second catalyst bed. The operating conditions for this hydrodesulfurization step in the second bed are as follows: lhsv=6h ^-1、H₂/HC＝262.5Nm³/m³.

Table 3 sets forth the characteristics of the liquid effluent produced in the second hydrodesulfurization step and the temperature conditions of the feed stream obtained therefrom over time.

TABLE 3 Table 3

Time of feed stream	Tiantian (Chinese character of 'Tian')	0	15	30	60	90
							Inlet sulfur content	ppmS	598	598	598	598	598
Outlet sulfur content	ppmS	10	10	11	12	13
							Bed 1 inlet temperature	℃	267	267	267	267	267
Bed 1 outlet temperature	℃	281	280	280	279	278
							Bed 2 inlet temperature	℃	276	276	275	274	273
Bed 2 outlet temperature	℃	286	285	285	283	281

* Olefin hydrogenation

Claims

1. A process for treating gasoline containing sulfur compounds, olefins and diolefins, said process comprising at least the steps of:

a) Contacting gasoline, hydrogen and a hydrodesulphurisation catalyst comprising an oxide support and an active phase comprising a group VIB metal and a group VIII metal at a temperature of 210 to 320 ℃, at a pressure of 1 to 4MPa, at a space velocity of 1 to 10h ^-1 and a ratio of the flow of hydrogen expressed as standard m ³/h to the flow of feedstock to be treated expressed as standard m ³/h of standard conditions of 100 to 600Nm ³/m³ to obtain a partially desulphurised effluent;

b) Directly contacting the partially desulphurised effluent obtained at the end of step a) with a gaseous or liquid feedstock as diluent under standard temperature and pressure conditions without any separation, to obtain a partially desulphurised and diluted effluent;

c) Contacting the partially desulphurised and diluted effluent obtained at the end of step b), hydrogen and a hydrodesulphurisation catalyst comprising an oxide support and an active phase comprising a group VIB metal and a group VIII metal, at a temperature of 210 to 320 ℃, at a pressure of 1 to 4MPa, at a space velocity of 1 to 10h ^-1, to obtain a desulphurised effluent.

2. The method of claim 1, wherein the diluent of step b) is a gaseous feedstock comprising at least 50% hydrogen by volume relative to the total volume of the gaseous feedstock.

3. The process according to claim 1 or 2, wherein the ratio of the volumetric flow rate of the gaseous feed introduced in step b) to the volumetric flow rate of hydrogen introduced in step a) is from 0.01 to 25Nm ³/Nm³.

4. The process of claim 1, wherein the diluent of step b) is a liquid hydrocarbon-based feedstock having a boiling point of from 30 ℃ to 250 ℃.

5. The process of claim 1 or 4 wherein the ratio of the volumetric flow rate of the liquid feedstock introduced in step b) to the volumetric flow rate of the gasoline feedstock introduced in step a) is from 0.01 to 1Sm ³/Sm³.

6. The process of claim 4 or 5 wherein the liquid feedstock introduced in step b) is the same as the gasoline feedstock introduced in step a) of the process.

7. The process according to any one of claims 1 to 6, wherein the gaseous feed or the liquid feed is introduced in step b) at a temperature lower than the temperature of the partially desulphurised effluent obtained at the end of step a).

8. The process of claim 7, wherein the gaseous feed or the liquid feed is introduced in step b) at a temperature of 20 ℃ to 300 ℃.

9. The process according to any one of claims 1 to 8, wherein steps a) and b) are carried out in the same reactor.

10. The process according to any one of claims 1 to 9, wherein steps a), b) and c) are carried out in the same reactor.

11. The process according to any one of claims 1 to 10, wherein the catalyst of step a) and/or step c) has a group VIII metal content of 0.1 to 10 wt% relative to the total weight of the catalyst and a group VIB metal content of 1 to 20 wt% relative to the total weight of the catalyst.

12. The process according to any one of claims 1 to 11, wherein the catalyst of step a) and/or step c) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, the catalyst comprising cobalt oxide in the form of CoO in an amount of from 0.1% to 10% relative to the total weight of the catalyst, molybdenum oxide in the form of MoO ₃ in an amount of from 1% to 20% relative to the total weight of the catalyst, the cobalt/molybdenum molar ratio being from 0.1 to 0.8, phosphorus oxide in the form of P ₂O₅ in an amount of from 0.3% to 10% relative to the total weight of the catalyst when phosphorus is present, the catalyst having a specific surface area of from 50 to 250m ²/g.

13. The process according to any one of claims 1 to 12, wherein the catalysts of steps a) and c) are the same.

14. The process according to any one of claims 1 to 13, wherein prior to step a) and prior to any optional distillation step, the gasoline is contacted with hydrogen and a selective hydrogenation catalyst to selectively hydrogenate the diolefins contained in the gasoline to give olefins.

15. The method of any one of claims 1 to 14, wherein the gasoline is a catalytically cracked gasoline.