Process for treating plastic and / or tire pyrolysis oil comprising removal of halide prior to hydrotreating step

Abstract

Disclosed is a process for treating tire and / or plastic pyrolysis oil comprising: a) hydrogenating said feedstock to obtain a hydrogenated effluent, b) separating the hydrogenated effluent at a temperature higher than the precipitation temperature of ammonium halide and under high pressure to obtain a gaseous effluent and a liquid effluent, c) hydrotreating the liquid effluent, and d) separating the hydrotreated effluent from the gaseous effluent. To obtain a hydrotreated effluent; d) feeding the gaseous effluent from steps b) and d) and the liquid effluent from step d) and the aqueous solution to the ammonium halide at a temperature higher than the precipitation temperature of the ammonium halide and under high pressure to obtain a gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

Description

Technical field

[0001] The present invention relates to a process for treating pyrolysis oil from plastics and / or tyres to obtain a hydrocarbon effluent which can be upgraded in a storage unit for gasoline, jet fuel or diesel fuel or as a feedstock for a steam cracking unit. More particularly, the present invention relates to a process for treating a feedstock obtained by pyrolysis of plastic waste or tyres between two catalytic stages using hydrogen to at least partially remove impurities and in particular halides. Prior art

[0002] Plastic waste is generally a mixture of several polymers, such as a mixture of polyethylene, polypropylene, polyethylene terephthalate, polyvinyl chloride and polystyrene. In addition, depending on the use, plastics may also contain other compounds in addition to the polymers, such as plasticisers, pigments, dyes or residues of polymerisation catalysts. Plastic waste may additionally contain, to a lesser extent, for example, biomass from household waste.

[0003] As regards tyres, they are mainly formed of rubber which provides elastic properties (a mixture of crosslinked synthetic and natural rubber type elastomers, with additives such as silica, resins, sulphur, zinc oxide, carbon black, etc.) and of fabrics and metal fibres which provide reinforcing properties.

[0004] Plastics or recycled tyres from collection and sorting channels may undergo a pyrolysis stage to obtain in particular pyrolysis oil. These oils generally contain many impurities, especially halogenated compounds, in particular chlorine-based compounds, but also contain dienes, olefins, metals, especially iron, silicon, or heteroelements such as sulphur, oxygen and nitrogen, and insoluble materials.

[0005] These pyrolysis oils from plastics and / or tyres are generally incinerated to produce electricity and / or used as fuel in industrial boilers or district heating boilers.

[0006] Another way of upgrading pyrolysis oil from plastics and / or tyres is to use these pyrolysis oils as a feedstock for a steam cracking unit to (re)produce olefins which are the constituent monomers of certain polymers.

[0007] Pyrolysis oil from plastics and / or tyres can also be upgraded to gasoline, jet fuel or diesel fuel.

[0008] However, pyrolysis oils from plastics and / or tires typically have a high content of impurities, and this content is incompatible with direct storage in a fuel storage unit or with a steam cracking unit or units located downstream of the steam cracking unit, in particular polymerization processes and selective hydrogenation processes. These impurities can cause operability problems and, in particular, problems of corrosion (especially due to the presence of chlorine), coking or catalytic deactivation, or incompatibility problems in the use of the target polymer. The presence of dienes can also lead to instability problems of the pyrolysis oil, characterized by the formation of gums. The gums and insoluble materials that may be present in the pyrolysis oil can cause plugging problems in the process.

[0009] One way to remove these impurities contained in plastics and / or tire pyrolysis oils is to perform hydrotreating (HDT) in the presence of a catalyst.

[0010] The most problematic impurities contained in plastics and / or tire oils are usually halides, and more particularly chlorine. This is because chlorine is generally a limiting contaminant in the treatment of pyrolysis oils in existing units of refineries. Chlorine is the cause (in the form of HCI) of corrosion that can occur in existing units whose metallurgy is generally not designed to tolerate even low chlorine levels. Another problem related to the presence of halides, especially chlorine, in pyrolysis oils is the formation of ammonium chloride salts, which form through the reaction between chloride ions (which are released as HCl by hydrodechlorination) and ammonium ions (which are generated as NH3 by the hydrogenation (hydrodenitrogenation) of nitrogen compounds during the hydrogenation and / or hydrotreating stages). It is known that these ammonium chloride salts precipitate at relatively low temperatures (e.g., below 280 °C), which causes plugging problems, especially in the transport lines and / or in process sections downstream of the hydrogenation / hydrotreating.

[0011] Methods for removing chlorine from pyrolysis oils by hydrotreating are known, for example, from the following documents: WO20020769, WO20016400, WO20239729, WO21105326 or WO16142809.

[0012] Document WO16142809 particularly describes a method that includes a pyrolysis stage, a hydrotreating stage, a separation stage of gases (C1 to C4 containing H2S and HCl) and C5+ liquids, a stage of dechlorinating the C5+ liquids by adsorption or a second hydrotreating stage, an optional separation stage of another HCl-containing gas, and then a steam cracking stage.

[0013] It is also known practice to remove HCl formed during hydrotreating by washing with water after the hydrotreating stage. Such washing is described in WO17083018, WO20254634 or WO22101333. None of these documents describe washing before the hydrotreating stage or the problem of deactivating the hydrotreating catalyst of a second hydrotreating stage downstream from the first hydrotreating stage due to the deposition of chlorinated compounds.

[0014] The unpublished patent application FR 21 / 12908 describes a process for treating plastic pyrolysis oil, which comprises:

[0015] a) a hydrogenation stage of a feedstock mixed with at least a portion of the liquid effluent from separation stage c) and hydrogen,

[0016] b) a hydrotreating stage of the effluent from stage a) in the presence of hydrogen,

[0017] c) a separation stage into which the effluent from stage b) is fed, said stage operating at a temperature between 200 and 450 °C and at a pressure substantially equal to that of stage b) to obtain at least a first gas effluent and a liquid effluent, with a portion of said liquid effluent being recycled upstream of stage a),

[0018] d) a separation stage into which the first gas effluent and another portion of the liquid effluent obtained from stage c) and an aqueous solution are fed, said stage operating at a temperature between 20 °C and less than 200 °C and at a pressure substantially equal to or less than that of stage c) to obtain at least a second gas effluent, an aqueous effluent and a hydrocarbon effluent.

[0019] One of the aims of application FR 21 / 12908 is to remove chlorine in the form of ammonium chloride salts by a combination of a thermal separation stage c) and a subsequent cold separation / washing stage d). Chloride ions released by the hydrogenation (hydrodechlorination) of chlorinated compounds in the form of HCl during stages a) and b) and ammonia formed by the hydrogenation (hydrodenitrogenation) of nitrogen-containing compounds in the form of NH3 especially during stage b) leave to a large extent in the gas effluent by virtue of the thermal separation of stage c). This is because the high temperature of this separation stage c) prevents the precipitation of ammonium chloride salts formed by the reaction between chloride ions and ammonium ions. The separation of the gas effluent and a portion of the liquid effluent at a lower temperature in stage d) causes the precipitation of these ammonium chloride salts. The water washing of this stage d) enables these salts to be dissolved in the aqueous effluent. A hydrocarbon effluent with chlorine removed is thus obtained.

[0020] The same principle of removing chlorine by thermal stripping to keep ammonia and halides in the gas phase until a certain amount of water is available to collect ammonium halides in solution is also described in WO22101333. In this document, as in application FR 21 / 12908, this separation is carried out after two HDT stages to send the purified oil to the steam cracking stage, aiming to avoid the precipitation of solid ammonium halides on the inner surface of the equipment and corrosion.

[0021] The present invention provides an improvement to the method according to FR 21 / 12908 by proposing to remove chlorine not after the hydrotreating stage b), but by performing a thermal separation before the hydrotreating stage b), in combination with a thermal separation and a cold separation and washing combination after the hydrotreating stage to remove chlorine.

[0022] The applicant has found that chlorine contained in the feedstock is almost completely converted during the hydrogenation stage. They have also observed that by directly feeding the effluent from the hydrogenation stage into the hydrotreating stage, chlorine (as HCl) contained in the effluent from the hydrogenation stage is deposited on the hydrotreating catalyst, thus resulting in a decrease in activity and a shortening of the duration of the catalytic cycle. To avoid clogging of the hydrotreating catalyst and a decrease in activity, the present invention proposes to remove chlorine before the hydrotreating stage.

[0023] Object and Summary of the Invention

[0024] More specifically, the present invention relates to a method for treating a feedstock comprising pyrolysis oil from plastics and / or tires containing halogen compounds, the method comprising:

[0025] a) a hydrogenation stage carried out in a hydrogenation reaction section using at least one fixed bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed comprising at least one hydrogenation catalyst, feeding to the hydrogenation reaction section at least the feedstock optionally as a mixture with at least a portion of a second liquid effluent obtained from a separation stage d) and a first gas stream containing hydrogen, the hydrogenation reaction section being used at an average temperature between 140 and 400 °C, a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a hydrogenation effluent,

[0026] b) a separation stage to which the hydrogenation effluent obtained from stage a) is fed, the stage operating at a temperature above the precipitation temperature of ammonium halides and at a pressure substantially equal to the pressure of stage a) to obtain at least a first gas effluent and a first liquid effluent,

[0027] c) A hydrotreating stage carried out in the hydrotreating reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed contains at least one hydrotreating catalyst, and at least the first liquid effluent obtained from stage b) and a second gas stream containing hydrogen are fed to the hydrotreating reaction section. The hydrotreating reaction section is used at an average temperature between 250 and 430 °C, a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a hydrotreated effluent,

[0028] d) A separation stage, to which the hydrotreated effluent obtained from stage c) is fed. The stage operates at a temperature higher than the precipitation temperature of the ammonium halide and at a pressure substantially equal to that of stage c) to obtain at least a second gas effluent and a second liquid effluent. Optionally, a part of the second liquid effluent is recycled to upstream of stage a).

[0029] e) A separation stage, to which the first and second gas effluents, at least a part of the second liquid effluent obtained from stage d), and an aqueous solution are fed. The stage operates at a temperature higher than the precipitation temperature of the ammonium halide and at a pressure substantially equal to or lower than that of stage d) to obtain at least a third gas effluent, an aqueous effluent, and a hydrocarbon effluent,

[0030] f) Optionally, a stage of fractionating all or part of the hydrocarbon effluent obtained from stage e) to obtain at least a fourth gas effluent, a naphtha fraction, and at least one middle distillate fraction,

[0031] g) Optionally, a hydrocracking stage carried out in the hydrocracking reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed contains at least one hydrocracking catalyst, and at least a part of the hydrotreated effluent obtained from stage c) and / or at least a part of the middle distillate fraction obtained from stage f) and a third gas stream containing hydrogen are fed to the hydrocracking reaction section. The hydrocracking reaction section is used at an average temperature between 250 and 450 °C, a hydrogen partial pressure between 1.5 and 20.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a first hydrocracked effluent.

[0032] The main advantage of the present invention is to separate the halogenated compounds contained in the hydrogenated effluent obtained from stage a) before the hydrotreating stage c) to protect the hydrotreating catalyst and at the same time avoid the precipitation and corrosion of solid ammonium halide on the inner surface of the equipment.

[0033] The halogenated compounds contained in the feedstock are almost completely converted into hydrogen halides (e.g., HCl) during the hydrogenation stage a), and some nitrogen compounds are converted into gaseous ammonia NH3. Therefore, in the thermal separation (stage b), i.e., at a temperature higher than the precipitation temperature of ammonium halide, ammonia and halides remain in the gas phase and a liquid effluent substantially free of halogen compounds is obtained, which is sent to a hydrotreating stage to remove the remaining impurities, and no deposition of halogen compounds on the catalyst is observed.

[0034] The hydrotreated effluent is then subjected again to a thermal separation stage (stage d)), i.e., at a temperature higher than the precipitation temperature of ammonium halide, which allows the pollutants (NH3 from HDN hydrodenitrogenation, H2S from HDS hydrodesulfurization, etc.) to remain in the gas phase and a hydrocarbon liquid effluent free of these pollutants is obtained.

[0035] The gas effluents from the first and second thermal separations (stages b) and d)) are then subjected to cold separation. In the cold separation (stage e), i.e., at a temperature lower than the precipitation temperature of ammonium halide and in the presence of an aqueous solution, these ammonium halide salts are precipitated while being dissolved in water.

[0036] Removing the halogen compounds before the hydrotreating stage allows the hydrotreating stage c) to operate at a lower average temperature than in the case where the halogen compounds are not removed beforehand, thereby saving energy.

[0037] In addition, an increase in the cycle time of the hydrotreating section is observed. The cycle time refers to the time during which the catalyst is used without the need to replace the catalyst.

[0038] In the first separation stage b) at a high temperature, almost all the light fractions of the hydrogenated effluent (naphtha fraction) leave as the first gas effluent and enter the cold separation / washing stage e), while the heavy fractions of the feedstock (middle distillate fraction) are mainly in the liquid phase. The heavy fraction is introduced as the first liquid effluent from stage b) into the hydrotreating stage c) for hydrotreating. The first gas effluent from stage b) is introduced into the cold separation stage e) and thus bypasses the hydrotreating stage c). By the method according to the present invention, the first gas effluent from stage b) is therefore not hydrotreated. Therefore, this first gas effluent may still contain pollutants (such as nitrogen, sulfur, etc.) that are conventionally removed during hydrotreating. The gas effluents (the first gas effluent from stage b) (untreated) and the second gas effluent from stage d) (treated)) are mixed with the second liquid effluent from stage d), which allows the use specifications of the pyrolysis oil in the upstream unit (such as a steam cracking unit) to be met by dilution. Therefore, the method according to the present invention is particularly suitable for "heavy" pyrolysis oil feedstocks, i.e., feedstocks mainly containing heavy fractions (middle distillate fractions).

[0039] The fact that the gas effluent from stage b) bypasses the hydrotreating stage c) allows the size of the reactor in the hydrotreating stage to be reduced, and thus the requirements for catalysts and equipment are saved.

[0040] The advantage of performing the cold separation stage e) on a mixture of gas effluent and liquid effluent (and not only on one or more gas effluents) is that it extracts pollutants not only from one or more gas effluents, but also hydrogen halide (HCl) dissolved in the hydrotreated liquid effluent.

[0041] Another major advantage of the present invention is that the halogen compounds of stage b) are separated at high pressure, and in particular at a pressure substantially the same as that of the hydrogenation stage a) (which is also substantially the same as the pressure of the hydrotreating stage c). Thus, the three stages a), b) and c) are carried out at the same pressure. Therefore, there is no need to reduce the pressure of the feedstock between stages a) and c), which represents an energy saving.

[0042] Another advantage of the method according to the present invention is to purify the oil obtained from the pyrolysis of plastic waste and / or tires by removing at least a part of its impurities, which allows it to be hydrogenated and thus upgraded, in particular by directly incorporating it into a fuel storage unit and / or by making it compatible with the treatment in a steam cracking unit, so as to be able to obtain, in particular with an increased yield, light olefins that can be used as monomers in polymer manufacturing.

[0043] Another advantage of the present invention is to prevent the risk of blockage and / or corrosion of the treatment unit in which the method of the present invention is carried out, which risk is exacerbated by the diolefins, metals and halogenated compounds often present in large amounts in the pyrolysis oil.

[0044] The method of the present invention thus makes it possible to obtain a hydrocarbon effluent from pyrolysis oil that is at least partially free of the impurities of the starting pyrolysis oil, thereby limiting the operability problems that these impurities may cause, such as corrosion, coking or catalytic deactivation problems, especially in a steam cracking unit and / or in a unit located downstream of the steam cracking unit, in particular polymerization and hydrogenation units. Removing at least a part of the impurities from the oil obtained from the pyrolysis of plastic waste will also make it possible to increase the scope of application of the target polymer to reduce the incompatibility of uses.

[0045] According to a variant, the method comprises a fractionation stage f).

[0046] According to a variant, the method comprises a hydrocracking stage g).

[0047] According to a variant, the method comprises a separation stage b') carried out between stages b) and c), into which at least a portion of the first liquid effluent obtained from stage b) and an aqueous solution are fed, the stage being operated at a temperature above the precipitation temperature of the ammonium halide and at a pressure substantially equal to or lower than the pressure of stage b), in order to obtain at least a gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

[0048] According to a variant, stage b) or stage d) is operated at a temperature between 200 and 450 °C, and stage e) or stage b') is operated at a temperature greater than or equal to 20 °C and less than 200 °C.

[0049] According to a variant, at least a portion of the second liquid effluent obtained in stage d) is recycled upstream of stage a).

[0050] According to a variant, at least a portion of the hydrocarbon effluent obtained from stage e) is recycled upstream of stage a) and / or upstream of stage c).

[0051] According to a variant, in stage a), the hydrogen coverage is between 250 and 800 Sm 3 hydrogen / m 3 feedstock (Sm 3 / m 3 ).

[0052] According to a variant, the method comprises at least one pre-treatment stage a0) of a fraction of the pyrolysis oil from plastics and / or tyres, the pre-treatment stage being carried out upstream of stage a) and comprising an adsorption stage and / or a filtration stage and / or a centrifugation stage and / or a sedimentation stage and / or an electrostatic separation stage and / or a stage of washing with an aqueous solution and / or a gas stripping stage.

[0053] According to a variant, at least one of the hydrocarbon effluent obtained from the separation stage e), or the fractions obtained from stage f), in whole or in part, is sent to a steam cracking stage h) carried out in at least one pyrolysis furnace at a temperature between 700 and 900 °C and a relative pressure between 0.05 and 0.3 MPa.

[0054] According to a variant, the reaction section of stage a) uses at least two reactors operating in a replaceable mode.

[0055] According to an alternative form, a stream containing an amine and / or a sulfur compound is injected upstream of stage a).

[0056] According to an alternative form, the hydrogenation catalyst comprises a support selected from alumina, silica, silica-alumina, magnesia, clays and mixtures thereof, and a hydrodehydrogenation functionality containing on the one hand at least one Group VIII element and at least one Group VIB element or on the other hand containing at least one Group VIII element.

[0057] According to a variant, the hydrotreating catalyst comprises a support selected from alumina, silica, silica-alumina, magnesia, clays and mixtures thereof, and a hydrodehydrogenation functionality containing at least one Group VIII element and / or at least one Group VIB element.

[0058] According to a variant, the method further comprises a second hydrocracking stage g') carried out in the hydrocracking reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed comprising at least one hydrocracking catalyst, feeding to the hydrocracking reaction section at least a portion of a first hydrocracking effluent obtained from the first hydrocracking stage g) and a gas stream containing hydrogen, the hydrocracking reaction section being used at a temperature between 250 and 450 °C, a hydrogen partial pressure between 1.5 and 20.0 MPa absolute pressure and a space velocity between 0.1 and 10.0 h -1 to obtain a second hydrocracking effluent.

[0059] According to a variant, the hydrocracking catalyst comprises a support selected from halogenated alumina, a combination of boron and aluminum oxides, amorphous silica-alumina and zeolites, and a hydrodehydrogenation functionality containing at least one Group VIB metal selected from chromium, molybdenum and tungsten (alone or as a mixture) and / or at least one Group VIII metal selected from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum.

[0060] The invention also relates to a product obtainable and preferably obtained by the method according to the invention.

[0061] According to this alternative form, the product comprises, relative to the total weight of the product:

[0062] - a total metal element content of less than or equal to 10.0 weight ppm,

[0063] - an iron element content of less than or equal to 200 weight ppb,

[0064] - a silicon element content of less than or equal to 5.0 weight ppm,

[0065] - a sulfur content of less than or equal to 500 weight ppm,

[0066] - a nitrogen content of less than or equal to 100 weight ppm,

[0067] - Chlorine element content less than or equal to 10 ppm by weight,

[0068] - Mercury content less than or equal to 5 ppb by weight.

[0069] According to the present invention, unless otherwise specified, the pressure is absolute pressure, also denoted as abs., and is given in MPa absolute pressure (or MPa abs.).

[0070] In this specification, the terms "comprising", "including" and "containing" are synonymous (have the same meaning), and are inclusive or open-ended, and do not exclude other elements that may not be mentioned. It is to be understood that the term "comprising" includes the exclusive and closed term "consisting of".

[0071] In this specification, unless otherwise specified, the expression "between... and... / ... to..." means that the limit values of the interval are included within the said numerical range.

[0072] Within the meaning of the present invention, the various parameter ranges of a given stage, such as the pressure range and the temperature range, can be used alone or in combination. For example, within the meaning of the present invention, a range of preferred pressure values can be combined with a range of more preferred temperature values.

[0073] Subsequently, specific and / or preferred embodiments of the present invention can be described. They can be implemented alone or in combination, and there is no limitation on the combination when technically feasible.

[0074] Hereinafter, unless otherwise specified, the term "pyrolysis oil" refers to an oil obtained by pyrolysis of plastics and / or tires.

[0075] Hereinafter, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, edited by D.R. Lide, 81st edition, 2000 - 2001). For example, Group VIII according to the CAS classification corresponds to the metals in columns 8, 9 and 10 according to the new IUPAC classification.

[0076] The metal content is measured by X-ray fluorescence.

[0077] Details

[0078] Raw materials

[0079] According to the present invention, "plastic pyrolysis oil" or "tyre pyrolysis oil" is advantageously an oil in liquid form at room temperature, which is obtained by pyrolysis of plastics, preferably pyrolysis of plastic waste, especially from collection and sorting channels, or from pyrolysis of used tyres.

[0080] It particularly includes mixtures of hydrocarbon compounds, especially alkanes, monoolefins and / or diolefins, naphthenes and aromatics. At least 80% by weight of these hydrocarbon compounds preferably have a boiling point below 700 °C, and preferably below 550 °C. In particular, depending on the source of the pyrolysis oil, the pyrolysis oil can contain up to 70% by weight of alkanes, up to 90% by weight of naphthenes, up to 90% by weight of olefins and up to 90% by weight of aromatics, it being understood that the sum of alkanes, naphthenes, olefins and aromatics is 100% by weight of the hydrocarbon compounds.

[0081] The oil can contain diolefins. The diolefin content is usually determined indirectly as the maleic anhydride value (MAV). This method is based on the Diels - Alder addition reaction between conjugated diolefins and maleic anhydride. The MAV determination method is described in C. López - García et al., Near Infrared Monitoring of Low Conjugated Diolefins Content in Hydrotreated FCC Gasoline Streams, Oil & Gas Science and Technology – Rev. IFP, Vol. 62 (2007), No. 1, pp. 57 - 68. The MAV is expressed as the number of milligrams of maleic anhydride reacting with 1 gram of the sample (mg / g). The MAV varies between 5 and 100 mg / g in the pyrolysis oil.

[0082] The density of the pyrolysis oil measured at 15 °C according to the ASTM D4052 method is usually between 0.75 and 0.99 g / cm 3 and preferably between 0.75 and 0.95 g / cm 3 between.

[0083] The pyrolysis oil can additionally contain, and usually does contain, impurities such as metals, especially iron, silicon or halogen compounds, especially chlorine compounds. These impurities can be present in the pyrolysis oil in high amounts, for example up to 600 weight ppm or 700 weight ppm, in fact even 1000 weight ppm, and even 5000 weight ppm of halogen elements (especially chlorine, but also bromine, fluorine and / or iodine) contributed by halogen compounds, and usually between 1 and 1000 weight ppm or between 1 and 700 weight ppm or between 1 and 600 weight ppm of halogen elements. The pyrolysis oil can contain up to 600 weight ppm or 700 weight ppm, or 1000 weight ppm and even 5000 weight ppm of chlorine element contributed by chlorinated compounds, and usually between 1 and 1000 weight ppm or between 1 and 700 weight ppm or between 1 and 600 weight ppm of chlorine element.

[0084] The oil may contain up to 200 weight ppm, and in fact even 1500 weight ppm, of metallic or metalloid elements, and generally between 1 and 200 weight ppm or between 1 and 1500 weight ppm of metallic or metalloid elements. Alkali metals, alkaline earth metals, transition metals, post-transition metals, and metalloids may be classified into the same category as contaminants with metallic properties, referred to as metallic or metalloid elements. In particular, the metallic or metalloid elements include silicon, iron, or both of these elements. The pyrolysis oil may particularly contain up to 200 weight ppm or 1000 weight ppm of silicon, and generally between 1 and 200 weight ppm or between 1 and 1000 weight ppm or between 1 and 500 weight ppm of silicon. The pyrolysis oil may particularly contain up to 50 weight ppm or 100 weight ppm of iron, and generally between 1 and 50 weight ppm or between 1 and 100 weight ppm of iron. The pyrolysis oil may also contain phosphorus, sodium, calcium, potassium, and magnesium.

[0085] The pyrolysis oil may also contain other impurities, such as heteroelements contributed particularly by sulfur compounds, oxygen compounds, and / or nitrogen compounds, with a content of generally less than 40000 weight ppm of heteroelements and preferably less than 15500 weight ppm of heteroelements, and generally between 1 and 40000 weight ppm or between 1 and 15500 weight ppm of heteroelements. Sulfur compounds are generally present in a content of less than 15000 weight ppm, and preferably less than 10000 weight ppm, and generally between 1 and 15000 weight ppm or between 1 and 10000 weight ppm of sulfur compounds.

[0086] Oxygen compounds are generally present in a content of less than 15000 weight ppm, and preferably less than 10000 weight ppm, and generally between 1 and 15000 weight ppm or between 1 and 10000 weight ppm of oxygen compounds.

[0087] Nitrogen compounds are generally present in a content of less than 10000 weight ppm, and preferably less than 5000 weight ppm, and generally between 1 and 10000 weight ppm or between 1 and 5000 weight ppm of nitrogen compounds.

[0088] The pyrolysis oil may also contain other impurities, such as heavy metals, such as mercury, arsenic, zinc, and lead, for example up to 100 weight ppb or 200 weight ppb of mercury or arsenic, and generally between 1 and 200 weight ppb or between 1 and 100 weight ppb of heavy metals.

[0089] The feedstock for the process according to the invention comprises at least one pyrolysis oil. The feedstock can consist solely of pyrolysis oil. Preferably, relative to the total weight of the feedstock, the feedstock comprises at least 50% by weight, preferably between 70% and 100% by weight, of pyrolysis oil, i.e. preferably between 50% and 100% by weight and preferably between 70% and 100% by weight of pyrolysis oil.

[0090] The feedstock for the process according to the invention may also comprise, in addition to the oil, conventional petroleum feedstocks or feedstocks obtained from biomass conversion, which are then co-processed with the pyrolysis oil of the feedstock.

[0091] The conventional petroleum feedstock may advantageously be a fraction or a mixture of fractions of the naphtha, gas oil or vacuum gas oil type.

[0092] The feedstock obtained from biomass conversion may advantageously be selected from vegetable oils, oils from algae or algal oils, fish oils, waste cooking oils and fats of plant or animal origin, or mixtures of these feedstocks. The oils / fats of plant and / or animal origin contain triglycerides and / or free fatty acids and / or esters. The vegetable oil may advantageously be crude or fully or partially refined and may be obtained from plants selected from rapeseed, sunflower, soybean, palm, olive, coconut, jatropha, copra, castor oil plant, cotton plant, peanut oil, linseed oil and sea kale oil, and all oils obtained, for example, from sunflower or rapeseed by genetic modification or hybridization, this list not being limiting. Algal oil or fish oil is also relevant. Animal fats are advantageously selected from lard and fats consisting of residues from the food industry or fats from the catering industry. Frying oils, various animal oils such as fish oil, tallow and lard may also be used.

[0093] The feedstock obtained from biomass conversion may also be selected from feedstocks derived from thermal or catalytic biomass conversion processes, such as oils produced from biomass, in particular lignocellulosic biomass, by various liquefaction methods, such as hydrothermal liquefaction or pyrolysis. The term "biomass" refers to materials derived from recently living organisms, which includes plants, animals and their by-products. The term "lignocellulosic biomass" refers to biomass derived from plants or their by-products. Lignocellulosic biomass consists of carbohydrate polymers (cellulose, hemicellulose) and aromatic polymers (lignin).

[0094] The feedstock obtained from biomass conversion may also advantageously be selected from feedstocks from the paper industry.

[0095] Plastic pyrolysis oil can be produced by thermal or catalytic pyrolysis treatment or also prepared by hydro-pyrolysis (pyrolysis in the presence of a catalyst and hydrogen).

[0096] (a0) Optional pretreatment stage

[0097] The raw material containing pyrolysis oil can be advantageously pretreated in an optional pretreatment stage a0) before the hydrogenation stage a), to obtain a pretreated raw material fed to stage a).

[0098] This optional pretreatment stage a0) makes it possible to reduce the amount of contaminants and solid particles that may be present in the raw material containing plastic pyrolysis oil, in particular the amount of iron and / or silicon and / or chlorine. Thus, especially when the raw material contains more than 10 ppm by weight, especially more than 20 ppm by weight, more especially more than 50 ppm by weight of metal elements and / or solid particles, and especially when the raw material contains more than 5 ppm by weight of silicon, more especially more than 10 ppm by weight, in fact even more than 20 ppm by weight of silicon, the optional pretreatment stage a0) of the raw material containing plastic pyrolysis oil is advantageously carried out. Similarly, especially when the raw material contains more than 10 ppm by weight, especially more than 20 ppm by weight, more especially more than 50 ppm by weight of chlorine, the optional pretreatment stage a0) of the raw material containing plastic pyrolysis oil is advantageously carried out.

[0099] The optional pretreatment stage a0) can be carried out by any method known to those skilled in the art that makes it possible to reduce the amount of contaminants. It can particularly include an adsorption stage and / or a filtration stage and / or a centrifugation stage and / or a sedimentation separation stage and / or an electrostatic separation stage and / or a stage of washing with an aqueous solution and / or a gas stripping stage.

[0100] This optional pretreatment stage a0) is advantageously carried out at a temperature between 20°C and 400°C, preferably between 40°C and 350°C and at an absolute pressure between 0.15 and 10.0 MPa, preferably between 0.2 and 7.0 MPa.

[0101] According to a variant, the optional pretreatment stage a0) is carried out in an adsorption section that operates in the presence of at least one adsorbent. The adsorbent can be selected from zeolites, activated carbon, clays, silica or alumina.

[0102] Advantageously, the adsorbent contains less than 1% by weight of metal elements and preferably is free of metal elements. The metal elements of the adsorbent should be understood to mean the elements from groups 6 to 10 of the periodic table (new IUPAC classification). The residence time of the raw material in the adsorption section is generally between 1 and 180 minutes.

[0103] The optional adsorption section of stage a0) comprises at least one adsorption tower, preferably at least two adsorption towers, preferably 2 to 4 adsorption towers, and the adsorption towers contain the adsorbent. When the adsorption section comprises two adsorption towers, one operating mode may be the "swing" operation according to the technical term, where one tower is online, i.e., in operation, while the other tower is on standby. When the adsorbent in the online tower is exhausted, the tower is isolated, and at the same time, the standby tower is put online, i.e., in operation. The exhausted adsorbent can then be regenerated in situ and / or replaced with fresh adsorbent so that once the other tower is isolated, the tower containing it can be put back online again.

[0104] Another operating mode is to operate at least two towers in series. When the adsorbent in the tower placed at the top is exhausted, the first tower is isolated, and the exhausted adsorbent is regenerated in situ or replaced with fresh adsorbent. The tower is then put back online at the last position, and so on. According to the technical term, this operation is called the permutable mode, or PRS, i.e., the permutable reactor system, or "lead and lag". The combination of at least two adsorption towers makes it possible to overcome the possible and potentially rapid poisoning and / or clogging of the adsorbent caused by the combined action of metal contaminants, diolefins, gums obtained from diolefins, and insolubles that may be present in the pyrolysis oil to be treated. This is because the presence of at least two adsorption towers promotes the replacement and / or regeneration of the adsorbent, advantageously without shutting down the pretreatment unit, or even the method in fact, and thus makes it possible to reduce the risk of clogging and thus avoid shutting down the unit due to clogging, in order to control costs and limit the consumption of the adsorbent.

[0105] According to another variant, the optional pretreatment stage a0) is carried out in a washing section using an aqueous solution, such as water or an acidic or alkaline solution. This washing section may comprise items of equipment capable of bringing the feedstock into contact with the aqueous solution and separating the phases, so as to obtain on the one hand the pretreated feedstock and on the other hand the aqueous solution containing impurities. These items of equipment may include, for example, stirred reactors, decanters, mixer-decanters, and / or co-current or counter-current washing towers.

[0106] According to another variant, the optional pretreatment stage a0) is carried out by filtration. The filtration stage makes it possible to remove the inorganic solids, deposits and / or fines contained in the oil, in particular metals, metal oxides and metal chlorides. Filters with a pore size (e.g., diameter or equivalent diameter) less than 25 μm, preferably less than or equal to 10 μm, and even more preferably less than or equal to 5 μm are generally used. According to another variant, filters with a pore size less than 25 μm but greater than 5 μm can be used. A series of filters with different pore sizes can also be used, in particular a series of filters with decreasing pore sizes in the oil flow direction. These filter media are well known for industrial use. Cartridge filters or self-cleaning filters are suitable, for example. The solids content can be measured, for example, by the heptane insolubles test, ASTM D-3279 method. The content of heptane insolubles must be reduced to less than 0.5% by weight, preferably reduced to less than 0.1%.

[0107] According to a specific embodiment, the pretreatment stage a0) carried out by filtration comprises at least one filter with a pore size less than 10 μm and preferably greater than 5 μm, optionally followed by a filtration system with a pore size less than 2 μm and preferably less than 1 μm.

[0108] According to another specific embodiment, the pretreatment stage a0) carried out by filtration comprises at least one filter with a pore size less than 10 μm and preferably greater than 5 μm, followed by an electrostatic precipitation system.

[0109] According to another specific embodiment, the pretreatment stage a0) carried out by filtration comprises at least one filter with a pore size less than 10 μm and preferably greater than 5 μm, followed by a filter system using a filter aid such as sand or diatomaceous earth.

[0110] According to another variant, the optional pretreatment stage a0) is carried out by centrifugation. According to another variant, the pretreatment stage a0) comprises centrifugation and filtration.

[0111] According to another variant, the optional pretreatment stage a0) is carried out by sedimentation. According to another variant, the pretreatment stage a0) comprises sedimentation and filtration.

[0112] According to another variant, said optional pretreatment stage a0) is carried out by gas stripping, thereby reducing the oxygen content in the oil. Gas extraction can remove the oxygen (O2) that may be dissolved in the feedstock, thereby reducing the possibility of the formation of free radicals that cause polymerization in downstream stages. This method generally involves contacting the oil with an extraction gas (such as H2, N2 or a mixture thereof), thereby transferring at least a part of the dissolved oxygen in the oil to the extraction gas, and then separating the extraction gas from the oil. The volume of the extraction gas relative to the oil volume (the two volumes measured under gas extraction conditions) is generally greater than 1, and preferably at least 3. In a specific embodiment, the extraction gas may contain at least 60% (mole percentage) of H2. Considering downstream hydrodemetallization / hydrotreatment, any dissolved H2 remaining in the feedstock after the gas extraction stage is not a problem. Preferably, the gas extraction stage is completed before any (pre)heating of the feedstock to minimize potential fouling.

[0113] Said optional pretreatment stage a0) generally comprises one or more, preferably a plurality of the above treatments. It may particularly comprise the following sequence: a washing stage and / or an adsorption stage using an aqueous solution, followed by a gas stripping stage, followed by a filtration stage and / or a centrifugation stage. All these stages are preferably carried out before any (pre)heating of the feedstock.

[0114] It is also possible to optionally feed at least a part of the liquid effluent from stage d) of the method and / or a part of at least one fraction from stage d) to said optional pretreatment stage a0), which is mixed with or separated from the feedstock containing pyrolysis oil. The recycling of at least a part of the liquid effluent from stage b) particularly enables an increase in sedimentation and thus an improvement in the pretreatment of the feedstock after possible filtration.

[0115] Said optional pretreatment stage a0) thus enables a pretreated feedstock to be obtained, which is subsequently fed to the hydrogenation stage a).

[0116] (a) Hydrogenation stage

[0117] According to the present invention, the method comprises a hydrogenation stage a) carried out in a hydrogenation reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed comprising at least one hydrogenation catalyst, feeding to said hydrogenation reaction section at least the feedstock optionally as a mixture with at least a part of a second liquid effluent obtained from the separation stage d) and a first gas stream containing hydrogen, said hydrogenation reaction section being used at an average temperature between 140 and 400 °C, a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure and a space velocity between 0.1 and 10.0 h -1 to obtain a hydrogenated effluent.

[0118] Stage a) is carried out in particular under hydrogen pressure and temperature conditions enabling the hydrogenation of dienes and olefins at the start of the hydrogenation reaction section, while allowing hydrodemetallation and hydrodechlorination to take place, in particular at the end of the hydrogenation reaction section, by means of an increasing temperature profile. The necessary amount of hydrogen is injected to allow the hydrogenation of at least a portion of the dienes and olefins present in the pyrolysis oil from the plastic, the hydrodemetallation of at least a portion of the metals, in particular the retention of silicon, and the conversion of at least a portion of the chlorine (to produce HCl). The hydrogenation of the dienes and olefins thus makes it possible to avoid or at least limit the formation of "muck", i.e. the polymerization of the dienes and olefins, and thus to limit the formation of oligomers and polymers which could clog the reaction section of the hydrotreating stage c). In parallel with the hydrogenation, hydrodemetallation takes place, and in particular the retention of silicon during stage a), which makes it possible to limit the catalytic deactivation of the reaction section of the hydrotreating stage c). Furthermore, the conditions of stage a) make it possible to convert at least a portion of the chlorine, and preferably all of the chlorine.

[0119] Those skilled in the art will readily understand that in the hydrogenation stage a), the hydrogenation reaction as described above is carried out, and also some other hydrotreating reactions are carried out in parallel, and in particular hydrodesulfurization and hydrodenitrogenation reactions, even though these reactions are rather advantageous in the hydrotreating stage c) which is usually carried out at a higher temperature.

[0120] The temperature in stage a), whether it is the average temperature (WABT), the inlet temperature of the reaction section or the temperature increase in stage a) between the inlet and the outlet of the reaction section, can be controlled in particular by the injection of a diluent in stage a), preferably by the recycle injection of a portion of the second liquid effluent obtained from stage d) and / or by the injection of at least a portion of one or more fractions obtained from stage g), in particular by the recycle rate and / or by the temperature of the recycle effluent.

[0121] The temperature difference between the inlet and the outlet of the reaction section of stage a) should be understood to refer to the injection of a gaseous (hydrogen) or liquid cooling stream, particularly preferably a portion of the second liquid effluent obtained from stage d).

[0122] The temperature difference between the inlet and the outlet of the reaction section of stage a) is entirely due to the exothermicity of the chemical reactions taking place in the reaction section and is thus compatible with the absence of heating devices (furnaces, heat exchangers, etc.).

[0123] The said reaction section is carried out in the presence of at least one hydrogenation catalyst, advantageously at an average temperature (or WABT as defined below) between 140 and 400 °C, preferably between 240 and 350 °C, and particularly preferably between 260 and 330 °C, at a hydrogen partial pressure between 1.0 and 10.0 MPa absolute, preferably between 1.5 and 8.0 MPa absolute and at 0.1 to 10.0 h -1between, preferably between 0.2 and 5.0 h -1 between, and very preferably between 0.3 and 3.0 h -1 The hydrogenation is carried out at a space velocity (HSV) between...

[0124] According to the invention, the "average temperature" of the reaction section corresponds to the weight average bed temperature (WABT) according to the specific terminology and is well known to those skilled in the art. The average temperature is advantageously determined according to the catalytic system used, the equipment and the configuration of the latter. The average temperature (or WABT) is calculated as follows:

[0125] WABT=(T entree +T sortie ) / 2

[0126] where T inlet : the temperature of the stream at the inlet of the reaction section, T outlet : the temperature of the effluent at the outlet of the reaction section. Unless otherwise stated, the "average temperature" of the reaction section is given under the conditions at the start of the cycle.

[0127] The space velocity (HSV) is defined herein as the ratio of the hourly volume flow rate of the feedstock containing pyrolysis oil, which has been optionally pre-treated, to the volume of the catalyst.

[0128] The hydrogen coverage is defined as the ratio of the hydrogen volume flow rate obtained under standard temperature and pressure conditions to the volume flow rate of the "fresh" feedstock (i.e., the feedstock to be treated that has been optionally pre-treated, without considering the recycle part, and in particular without considering the recycle liquid effluent obtained from stage d) at 15 °C (expressed as Sm 3 (denoted as Sm 3 ) of H2 / m 3 feedstock).

[0129] The amount of the gas stream containing hydrogen (H2) fed to the reaction section of stage a) is advantageously such that the hydrogen coverage is between 100 and 1500 Sm 3 hydrogen / m 3 feedstock (Sm 3 / m 3 ), preferably between 200 and 1000 Sm 3 hydrogen / m 3 feedstock (Sm 3 / m 3 ), and in a preferred manner between 250 and 800 Sm 3 hydrogen / m 3 feedstock (Sm 3 / m 3 ).

[0130] Advantageously, the reaction section of stage a) comprises from 1 to 5 reactors, preferably from 2 to 5 reactors, and particularly preferably comprises two reactors. The advantage of a hydrogenation reaction section comprising a plurality of reactors lies in optimizing the treatment of the raw materials, while making it possible to reduce the risk of clogging of the catalytic bed and thus avoid the shutdown of the unit caused by clogging.

[0131] According to a preferred variant, these reactors operate in a replaceable mode, called PRS, i.e., a replaceable reactor system, or "lead and lag". The combination of at least two reactors in PRS mode makes it possible to isolate one reactor to discharge the spent catalyst, refill the reactor with fresh catalyst and put the reactor back into use without shutting down the process. The PRS technique is described in particular in patent FR2681871.

[0132] According to a particularly preferred variant, the hydrogenation reaction section of stage a) comprises two reactors operating in a replaceable mode.

[0133] Advantageously, reactor internals, such as reactor internals of the filter plate type, can be used to prevent clogging of the reactor. An example of a filter plate is described in patent FR3051375.

[0134] Advantageously, the hydrogenation catalyst comprises a support, preferably an inorganic support, and a hydrodehydrogenation functionality.

[0135] According to a variant, the hydrodehydrogenation functionality particularly comprises at least one Group VIII element preferably selected from nickel and cobalt, and at least one Group VIB element preferably selected from molybdenum and tungsten. According to this variant, the total content of Group VIB and VIII metal elements expressed as oxides is preferably between 1% by weight and 40% by weight, preferably between 5% by weight and 30% by weight, relative to the total weight of the catalyst. When the metal is cobalt or nickel, the metal content is expressed as CoO and NiO respectively. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 and WO3 respectively.

[0136] The weight ratio of the Group VIB metal expressed as metal oxide to the Group VIII metal is preferably between 1 and 20, and preferably between 2 and 10.

[0137] According to this variant, the reaction section of stage a) comprises a hydrogenation catalyst comprising, for example, 0.5% by weight to 12% by weight of nickel, preferably 0.9% by weight to 10% by weight of nickel (expressed as nickel oxide NiO relative to the weight of the catalyst), and 1% by weight to 30% by weight of molybdenum, preferably 3% by weight to 20% by weight of molybdenum (expressed as molybdenum oxide MoO3 relative to the weight of the catalyst), on a preferably inorganic support, preferably an alumina support.

[0138] According to another variant, the hydrogenation-dehydrogenation functionality comprises at least one Group VIII element, preferably nickel, and is preferably composed of at least one Group VIII element, preferably nickel. According to this variant, the nickel oxide content is preferably between 1% by weight and 50% by weight, preferably between 10% by weight and 30% by weight, based on the weight of the catalyst. This type of catalyst is preferably used in its reduced form, on an inorganic support, preferably on an alumina support.

[0139] The support of the hydrogenation catalyst is preferably selected from alumina, silica, silica-alumina, magnesia, clay and mixtures thereof. The support may comprise dopant compounds, in particular oxides selected from boron oxide, more particularly boron trioxide, zirconium dioxide, cerium dioxide, titanium oxide, phosphorus pentoxide and mixtures of these oxides. Preferably, the hydrogenation catalyst comprises an alumina support optionally doped with phosphorus and optionally doped with boron. When phosphorus pentoxide P2O5 is present, its concentration is less than 10% by weight based on the weight of the alumina, and advantageously at least 0.001% by weight based on the total weight of the alumina. When boron trioxide B2O3 is present, its concentration is less than 10% by weight based on the weight of the alumina, and advantageously at least 0.001% by weight based on the total weight of the alumina. The alumina used may be, for example, γ (gamma) or η (eta) alumina.

[0140] The hydrogenation catalyst is in the form of, for example, an extrudate.

[0141] Very preferably, in stage a), in addition to the above hydrogenation catalyst, at least one hydrogenation catalyst used in stage a) can be used, which comprises less than 1% by weight of nickel, expressed as nickel oxide NiO, and at least 0.1% by weight of nickel, preferably 0.5% by weight of nickel, and less than 5% by weight of molybdenum, expressed as molybdenum trioxide MoO3, and at least 0.1% by weight of molybdenum, preferably 0.5% by weight of molybdenum, based on the weight of the catalyst, on an alumina support. This catalyst with a low metal loading can preferably be arranged upstream or downstream of the above hydrogenation catalyst, preferably upstream.

[0142] Preferably, in stage a), at least one guard bed can be used upstream of the hydrogenation catalyst, which contains adsorbents of the alumina, silica-alumina, zeolite and / or activated carbon type, optionally containing metals of Group VIB and / or Group VIII. A series of guard beds with particles of different diameters can also be used, especially a series of guard beds with decreasing diameters in the direction of the raw material flow (also called "graded").

[0143] According to one variant, the feedstock can be preheated by direct heating to a temperature of at most 200 °C, preferably at most 180 °C, and particularly preferably at most 150 °C, before being preferably mixed with at least a portion of the second effluent obtained from stage d). Above this temperature, contact with the walls during direct heating can lead to the formation of sludge and / or coke, which can cause fouling of the feedstock heating system and the catalyst bed and an increase in pressure drop. The feedstock is preferably heated to a temperature above 150 °C, preferably above 180 °C, and particularly preferably above 200 °C by indirect heating with a hot diluent, preferably with at least a portion of the second effluent obtained from stage d). Thus, the temperature of the feedstock is increased to above 150 °C, preferably above 180 °C, and particularly preferably above 200 °C by mixing with a hotter liquid rather than by contact with the heating walls. This makes it possible to locally limit high temperatures. This type of heating by mixing with an inert hot liquid thus makes it possible to limit undesired reactions such as the polymerization of diolefins (sludge formation) and / or the formation of coke, and to adjust the inlet temperature of the stream in stage a) to preferably initiate the hydrogenation reaction of unsaturations at the lowest possible temperature, while controlling the exothermicity of these reactions by the dilution effect of the reactive entities.

[0144] According to another variant, the feedstock is heated entirely by indirect heating with at least a portion of the second effluent obtained from stage d). In this case, the feedstock is not preheated before being mixed with at least a portion of the second effluent obtained from stage d).

[0145] Thus, the energy required for the reaction is advantageously achieved by mixing the feedstock containing pyrolysis oil and the hydrogen-rich gas with a diluent and preferably with a portion of the second liquid effluent obtained from the separation stage d), which has optionally been temperature-regulated and preferably preheated or cooled and particularly preferably preheated, upstream of stage a), and more specifically, the adjustment of the minimum temperature required to activate the double bond saturation reaction.

[0146] Another heating stream advantageously consists of a hydrogen-rich gas effluent originating from the hydrogen supply and / or a gas effluent obtained from separation stage e). At least a portion of this hydrogen-rich gas effluent originating from the hydrogen supply and / or at least a portion of the gas effluent obtained from separation stage e) is advantageously injected upstream of stage a) either as a mixture with at least a portion of the second effluent obtained from stage d) or separately.

[0147] The hydrogenation step b) makes it possible to obtain a hydrogenated effluent, i.e. an effluent having a reduced content of olefins (especially diolefins), metals (especially silicon) and halogens (especially chlorine). The impurity content of the hydrogenated effluent obtained at the end of step a), especially the diolefin content, is reduced relative to the content of the same impurities, especially the diolefin content, contained in the feedstock of the process. The hydrogenation step a) generally makes it possible to convert at least 40%, and preferably at least 60%, of the diolefins contained in the initial feedstock and at least 40% and preferably at least 60% of the olefins. The heat released by the saturation of the double bonds makes it possible to increase the temperature of the reaction medium and initiate the hydrotreating reaction, especially to remove at least partially other contaminants such as, for example, silicon and chlorine, or nitrogen. Preferably, during step a), at least 50%, and more preferably at least 75%, of the chlorine and silicon of the initial feedstock are removed respectively. Generally, the silicon content is less than 10 ppm by weight. The effluent obtained at the end of the hydrogenation step a) is preferably sent directly to the thermal separation step b).

[0148] (b) First thermal separation step

[0149] According to the invention, the treatment process comprises a separation step b) which is fed with the hydrogenated effluent obtained from step a), said step being carried out at a temperature above the precipitation temperature of the ammonium halide and at a pressure substantially equal to that of step a) to obtain at least a first gaseous effluent and a first liquid effluent.

[0150] This elevated temperature of the separation step b) avoids the precipitation of ammonium halide salts formed by the reaction between halide ions and ammonium ions in order to recover a first gaseous effluent containing most of the halides in the form of hydrogen halide (HCl) and gaseous ammonia, and a first liquid effluent containing very low amounts of halides and ammonia.

[0151] The "precipitation temperature" of the ammonium halide is understood to mean the temperature at which gaseous ammonia and hydrogen halide react (under given conditions such as concentration and pressure) to form solid crystals of ammonium halide or precipitate by dissolving in water. According to the principles of thermodynamics, the precipitation temperature depends on the concentration of the halide and the pressure. Under the operating conditions of the process, the precipitation temperature of the ammonium halide is generally between 150 and 300 °C, often between 180 and 295 °C, and especially between 200 and 290 °C.

[0152] The temperature at which the separation in step b) is carried out must be higher than the precipitation temperature of the ammonium halide in order to recover a first gaseous effluent containing most of the halides in the form of hydrogen halide (HCl) and ammonia gas, and a first liquid effluent containing very low amounts of halides and ammonia. The temperature at which the separation in step b) is carried out is generally between 200 and 450 °C, preferably between 220 and 330 °C, and particularly preferably between 240 and 300 °C.

[0153] The term "pressure substantially equal to the pressure of stage a)" is understood to mean a pressure difference relative to the pressure of stage a) that is between 0 and 1 MPa, preferably between 0.005 and 0.3 MPa, and particularly preferably between 0.01 and 0.3 MPa of the pressure of stage a). Preferably, the pressure of stage b) is the pressure of stage a) minus the pressure drop.

[0154] The separation stage b) can advantageously be carried out by any method known to the person skilled in the art, such as, for example, a combination of one or more separators (drums) and / or one or more stripping columns, which separator(s) (drum) and / or column(s) can optionally be fed with stripping gas, such as a hydrogen-rich gas stream. Preferably, stage b) is carried out with a single separator (drum).

[0155] The separation stage b) is a separation stage known as a high-pressure or medium-pressure high-temperature separation stage and is also known to the person skilled in the art by the name HHPS (i.e., hot high-pressure separator). Thus, this stage b) preferably uses a "hot high-pressure" separator with a pressure substantially equal to the operating pressure of stage a).

[0156] This gas / liquid separation has an efficiency corresponding to solubility and Henry's law. This means that an equilibrium amount of the halide in the form of hydrogen halide will remain in the first liquid effluent obtained in stage b). These halides will be released into the gas effluent during the cold separation stage e) and then discharged by dissolving in the aqueous effluent.

[0157] The thermal separation of stage b) allows ammonia and halides to remain in the gas phase and obtains a first liquid effluent substantially free of halogen compounds, which can be sent to the hydrotreating stage c) to remove the remaining impurities without observing the deposition of halogen compounds on the catalyst.

[0158] According to the first variant, the first liquid effluent is sent in part or in whole (and preferably in whole) to the hydrotreating stage c).

[0159] According to another variant, the first liquid effluent can be sent to a cold separation / additional washing stage b') before it is sent to the hydrotreating stage c).

[0160] In this case, the method according to the invention comprises a separation stage b') which is fed with at least a part of the first liquid effluent from stage b) and an aqueous solution, said stage operating at a temperature below the precipitation temperature of ammonium halide and at a pressure substantially equal to or below the pressure of stage b) to obtain at least a gas effluent, an aqueous effluent and a hydrocarbon effluent. In this way, the halides in the form of hydrogen halide dissolved in the first liquid effluent are released into the gas effluent during the cold separation stage b') and subsequently dissolve in the aqueous effluent.

[0161] This cold separation / washing separation stage b') operates in substantially the same manner as separation stage e) described below, and the separation conditions (pressure and temperature) can be the same or different.

[0162] (c) Hydrotreating stage

[0163] According to the present invention, the treatment method comprises a hydrotreating stage d) carried out in a hydrotreating reaction section using at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed comprising at least one hydrotreating catalyst, feeding at least the first liquid effluent obtained from stage b) and a second gas stream containing hydrogen to the hydrotreating reaction section, the hydrotreating reaction section being used at an average temperature between 250 and 430 °C, a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a hydrotreated effluent.

[0164] Advantageously, stage c) implements hydrotreating reactions well-known to those skilled in the art, and more particularly hydrotreating reactions such as the hydrogenation of aromatics, hydrodesulfurization, and hydrodenitrogenation. In addition, the hydrogenation of remaining olefins and halogen compounds and hydrodemetallation can be continued, even if most, and preferably all, of these impurities have been removed during stage a).

[0165] The hydrotreating reaction section is advantageously used at a pressure equal to that used in the reaction section of hydrogenation stage a) and generally at an average temperature higher than that of the reaction section of hydrogenation stage a). Thus, the hydrotreating reaction section is advantageously used at an average hydrotreating temperature between 250 and 430 °C, preferably between 280 and 380 °C, at a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure, and at a space velocity (HSV) between 0.1 and 10.0 h -1 preferably between 0.1 and 5.0 h -1 preferably between 0.2 and 2.0 h -1 preferably between 0.2 and 1 h -1 preferably between 0.2 and 1 h 3 hydrogen / m 3 fresh feedstock fed to stage a), and preferably between 200 and 1000 Sm 3 hydrogen / m 3 fresh feedstock fed to stage a), preferably between 250 and 800 Sm 3 hydrogen / m 3 fresh feedstock fed to stage a). The definitions of average temperature (WABT), HSV, and hydrogen coverage correspond to those described above.

[0166] At least the first liquid effluent obtained from stage b) and a second gas stream comprising hydrogen are fed to the hydrotreating reaction stage, advantageously at the level of the first catalytic bed of the reactor in the first run.

[0167] Advantageously, stage c) is carried out in a hydrotreating reaction stage comprising at least one, preferably 1 to 5, fixed-bed reactors having n catalytic beds, where n is an integer greater than or equal to 1, preferably 1 to 10, preferably between 2 and 5, each bed comprising at least one, and preferably no more than 10, hydrotreating catalysts. When the reactor comprises a plurality of catalytic beds, i.e., at least 2, preferably 2 to 10, preferably 2 to 5, catalytic beds, the catalytic beds are preferably arranged in series in the reactor.

[0168] When stage c) is carried out in a hydrotreating reaction stage comprising a plurality of reactors, preferably two reactors, these reactors can be operated in series and / or in parallel and / or in a replaceable (or PRS) mode and / or in a swing mode. The various possible operating modes, PRS mode (or lead or lag) mode, and swing mode are well known to those skilled in the art and are advantageously defined above.

[0169] In another embodiment of the present invention, the hydrotreating reaction stage comprises a single fixed-bed reactor containing n catalytic beds, where n is an integer greater than or equal to 1, preferably 1 to 10, preferably 2 to 5.

[0170] In a particularly preferred embodiment, the hydrogenation reaction stage of stage a) comprises two reactors operated in a replaceable mode, followed by a hydrotreating reaction stage of stage c) comprising a single fixed-bed reactor.

[0171] Advantageously, the hydrotreating catalysts used in stage c) can be selected from known hydrodemetallation, hydrotreating, and / or silicon capture catalysts particularly used for treating petroleum fractions, and combinations thereof. Known hydrodemetallation catalysts are, for example, those described in patents EP 0113297, EP 0113284, US 5221656, US 5827421, US 7119045, US 5622616, and US5089463. Known hydrotreating catalysts are, for example, those described in patents EP 0113297, EP 0113284, US6589908, US 4818743, or US 6332976. Known silicon capture catalysts are, for example, those described in patent applications CN102051202 and US2007 / 080099.

[0172] In particular, the hydrotreating catalyst comprises a support, preferably an inorganic support, and at least one metal element having hydrodehydrogenation functionality. The metal element having hydrodehydrogenation functionality advantageously comprises at least one Group VIII element preferably selected from nickel and cobalt, and / or at least one Group VIB element preferably selected from molybdenum and tungsten. Based on the total weight of the catalyst, the total content of Group VIB and Group VIII metal elements expressed as oxides is preferably between 0.1% by weight and 40% by weight, preferably 5% by weight and 35% by weight. When the metal is cobalt or nickel, the metal content is expressed as CoO and NiO respectively. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 and WO3 respectively. The weight ratio of the Group VIB metal to the Group VIII metal expressed as metal oxides is preferably between 1.0 and 20, and preferably between 2.0 and 10. For example, the hydrotreating reaction section in stage d) of the process comprises a hydrotreating catalyst comprising 0.5% by weight to 10% by weight of nickel expressed as nickel oxide NiO, preferably 1% by weight to 8% by weight of nickel, and 1.0% by weight to 30% by weight of molybdenum expressed as molybdenum oxide MoO3, preferably 3.0% by weight to 29% by weight of molybdenum, based on the total weight of the hydrotreating catalyst, on an inorganic support, preferably on an alumina support.

[0173] The support of the hydrotreating catalyst is advantageously selected from alumina, silica, silica-alumina, magnesia, clays and mixtures thereof. The support may additionally include dopant compounds, in particular oxides selected from boron oxide, especially boric anhydride, zirconia, ceria, titania, phosphorus pentoxide and mixtures of these oxides. Preferably, the hydrotreating catalyst comprises an alumina support, preferably an alumina support doped with phosphorus and optionally boron. When phosphorus pentoxide P2O5 is present, its concentration is less than 10% by weight relative to the weight of alumina, and advantageously at least 0.001% by weight relative to the total weight of alumina. When boric anhydride B2O3 is present, its concentration is less than 10% by weight relative to the weight of alumina, and advantageously at least 0.001% by weight relative to the total weight of alumina. The alumina used may be, for example, γ (gamma) or η (eta) alumina.

[0174] The hydrotreating catalyst is in the form of, for example, an extrudate.

[0175] Advantageously, the hydrotreating catalyst of the process has a specific surface area greater than or equal to 250 m 2 / g, preferably greater than or equal to 300 m 2 / g. The specific surface area of the hydrotreating catalyst is advantageously less than or equal to 800 m 2 / g, preferably less than or equal to 600 m 2 / g, particularly less than or equal to 400 m 2 / g. The specific surface area of the hydrotreating catalyst is measured by the BET method, i.e., the specific surface area determined by nitrogen adsorption according to the standard ASTM D 3663-78 established by the Brunauer-Emmett-Teller method described in The Journal of the American Chemical Society, 6Q, 309 (1938). Such a specific surface area enables further improvement in the removal of contaminants, particularly metals such as silicon.

[0176] According to another aspect of the present invention, the hydrotreating catalyst as described above further comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur. Such catalysts are generally denoted by the term "additivated catalyst". Generally, the organic compound is selected from compounds containing one or more chemical functional groups selected from carboxyl, alcohol, thiol, thioether, sulfone, sulfoxide, ether, aldehyde, ketone, ester, carbonate, amine, nitrile, imide, oxime, urea, and amide functional groups, or compounds containing a furan ring or sugar.

[0177] Preferably, stage c) can use, upstream of the hydrogenation catalyst, at least one guard bed or a series of guard beds of the "graded" type as described above for stage a).

[0178] Advantageously, the hydrotreating stage c) allows the hydrogenation of at least 80%, and preferably all, of the possible olefins remaining after the hydrogenation stage a) and the possible halogen compounds remaining in the first liquid effluent from the separation stage b), and also allows at least partial conversion of other impurities present in the feedstock, such as aromatic compounds, metal compounds, sulfur compounds, nitrogen compounds, and / or oxygen compounds. Preferably, the nitrogen content at the outlet of stage c) is less than 100 weight ppm and preferably less than 10 weight ppm. Preferably, the sulfur content at the outlet of stage c) is less than 100 weight ppm and preferably less than 10 weight ppm. Stage c) can also enable further reduction in the content of contaminants, such as the content of metals, particularly the silicon content. Preferably, the metal content at the outlet of stage c) is less than 10 weight ppm and preferably less than 2 weight ppm, and the silicon content is less than 5 weight ppm. Preferably, the content of halogen elements at the outlet of stage c) is less than 5 weight ppm.

[0179] Depending on the content of sulfur compounds in the initial feedstock to be treated, a stream containing a sulfurizing agent can be injected upstream of the hydrogenation stage a) and / or the hydrotreating stage c) and / or upstream of one of the hydrocracking stages (when they are present), preferably upstream of the hydrogenation stage a) and / or the hydrotreating stage c), to ensure that a sufficient amount of sulfur is present to form or maintain the active entities of the catalyst (in sulfided form). This activation or sulfidation stage is carried out by methods known to those skilled in the art and is advantageously carried out in a sulfo-reductive atmosphere in the presence of hydrogen and hydrogen sulfide. The sulfurizing agent is preferably H2S gas, elemental sulfur, CS2, mercaptans, sulfides and / or polysulfides, hydrocarbon fractions containing sulfur compounds with a boiling point less than 400 °C, or any other sulfur-containing compound used to activate the hydrocarbon feedstock for sulfiding the catalyst. The sulfur-containing compound is advantageously selected from alkyl disulfides such as, for example, dimethyldisulfide (DMDS), alkyl sulfides such as, for example, dimethyl sulfide, mercaptans such as, for example, n-butyl mercaptan (or 1-butyl mercaptan), and polysulfide compounds of the tert-nonyl polysulfide type. The catalyst can also be sulfided with the sulfur contained in the feedstock to be desulfurized. Preferably, the catalyst is sulfided in situ in the presence of a sulfurizing agent and a hydrocarbon feedstock. Very preferably, the catalyst is sulfided in situ in the presence of a feedstock to which dimethyldisulfide has been added.

[0180] (d) Second thermal separation stage

[0181] According to the invention, the treatment method comprises a separation stage d) to which is fed the hydrotreating effluent obtained from stage c). The stage operates at a temperature above the precipitation temperature of ammonium halide and at a pressure substantially the same as that of stage c) to obtain at least a second gas effluent and a second liquid effluent, a portion of the second liquid effluent being optionally recycled upstream of stage a).

[0182] This second thermal separation stage d) operates in substantially the same manner as the first separation stage b), and the separation conditions (pressure and temperature) can be the same or different.

[0183] Although the vast majority of halogen compounds are removed during the first thermal separation stage b), the high temperature of this separation stage d) avoids the precipitation of ammonium halide salts, which are formed by the reaction of halide ions with ammonium ions formed during the hydrotreating of the remaining halogen compounds. The high temperature also makes it possible to remove the NH3 and H2S formed respectively during the hydrotreating of nitrogen and sulfur compounds by the second gas effluent.

[0184] The temperature at which the separation of stage d) is carried out is generally between 200 and 450 °C, preferably between 220 and 330 °C, and particularly preferably between 240 and 300 °C.

[0185] The term "pressure substantially equal to the pressure of stage c)" shall be understood to mean that the pressure differential of the pressure of stage c) relative to the pressure of stage c) is between 0 and 1 MPa, preferably between 0.005 and 0.3 MPa, and particularly preferably between 0.01 and 0.3 MPa. Preferably, the pressure of stage d) is the pressure of stage c) minus the pressure drop.

[0186] The separation stage d) can advantageously be carried out by any method known to those skilled in the art, such as, for example, a combination of one or more separators (drums) and / or one or more stripper columns, and a stripping gas, such as a hydrogen-rich feed stream, can optionally be fed to this or these separators (drums) and / or columns. Preferably, stage d) is carried out using a single separator (drum).

[0187] The separation stage d) is a separation stage known as a high-pressure or medium-pressure high-temperature separation stage, and is also known to those skilled in the art by the name HHPS (hot high-pressure separator). Therefore, stage d) preferably employs a "hot high-pressure" separator, the pressure of which is substantially equal to the operating pressure of stage c).

[0188] This gas / liquid separation has an efficiency corresponding to solubility and Henry's law. This means that an equilibrium amount of the halide in the form of hydrogen halide will remain in the second liquid effluent obtained in stage b). These halides will be released into the gas effluent during the cold separation stage e) and will subsequently be discharged by dissolving in the aqueous effluent.

[0189] Advantageously, part of the second liquid effluent from stage d) is recycled upstream of stage a).

[0190] Upstream of stage a), the feedstock is mixed with the second liquid effluent (hot recycle), so that on the one hand, the impurities in the feedstock can be diluted, and on the other hand, the temperature in stage a) can be controlled, in which a highly exothermic reaction (especially the hydrogenation of olefins and diolefins) occurs. Upstream of stage a), mixing the feedstock with part of the second liquid effluent (hot recycle) also allows the feedstock to be indirectly heated simply by mixing the "cold" feedstock with the hot recycle.

[0191] Adjust the amount of the recycled second liquid effluent from stage d) such that the weight ratio between the recycle stream from stage d) and the feedstock containing pyrolysis oil (i.e., the feedstock to be treated in the overall process) is less than or equal to 10, preferably less than or equal to 7, and more preferably greater than or equal to 0.001, preferably greater than or equal to 0.01, and more preferably greater than or equal to 0.1. Preferably, adjust the amount of the recycled second liquid effluent from stage d) such that the weight ratio between the recycle stream and the feedstock containing plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5. This recycle rate enables the control of the temperature rise in stage a). This is because when the recycle rate is high, the dilution rate of the feedstock is high, and especially due to the hydrogenation reaction of dienes, the temperature rise at the start of the reaction section in stage a) can thus be controlled by the dilution effect.

[0192] According to a preferred variant, at least a portion of the second liquid effluent obtained from stage d) can be advantageously alternately cooled, or preheated if necessary, or maintained at the same temperature as at the outlet of the separation stage d), depending on the temperature and flow rate of the feedstock and hydrogen, prior to preferably recycling upstream of the hydrogenation stage a), such that the temperature of the incoming feed stream containing the feedstock (preferably as a mixture with at least a portion of the second liquid effluent obtained from stage d) and the hydrogen-rich gas) is between 140 and 400 °C, preferably between 220 and 350 °C, and particularly preferably between 260 and 330 °C.

[0193] High-pressure and high-temperature separation particularly enables the maximization of energy recovery by thermally recycling a portion of the second liquid effluent. This is because the energy required to reach the inlet temperature for stage a) is at least partially contributed by the heat of at least a portion of the second liquid effluent obtained from stage d), and also enables the reduction or even elimination of optional preheating (by directly heating the feedstock to a temperature above 200 °C) to prevent sludge formation. Furthermore, it is preferred to recycle at least a portion of the second liquid effluent under high pressure, such that the energy used for its pressurization in stage a) can be saved.

[0194] (e) Cold separation stage and washing

[0195] According to the invention, the treatment method includes a separation stage e), to which the first and second gas effluents, at least a portion of the second liquid effluent obtained from stage d), and an aqueous solution are fed. The stage operates at a temperature below the precipitation temperature of ammonium halide and at a pressure substantially equal to or below the pressure of stage d) to obtain at least a third gas effluent, an aqueous effluent, and a hydrocarbon effluent.

[0196] In stage e), at a temperature below the precipitation temperature of the ammonium halide, a mixture of the gaseous effluent and at least a portion of the second liquid effluent from stage d) is separated, causing these ammonium halide salts to precipitate. In addition, the hydrogen halide dissolved in the second liquid effluent from stage d) is released and also forms ammonium halide with the ammonia present in the gaseous effluent. Washing with an aqueous solution in this stage e) allows these salts to dissolve in the aqueous effluent. A hydrocarbon effluent free of halides, a gaseous effluent free of halides, and an aqueous effluent in which the ammonium halide salts are dissolved are thus obtained.

[0197] The temperature at which the separation in stage e) is carried out should be below the precipitation temperature of the ammonium halide to cause the ammonium halide salts to precipitate. The temperature at which the separation in stage e) is carried out is between 20 and below 200 °C, preferably between 25 and 120 °C, and particularly preferably between 30 and 70 °C.

[0198] The term "pressure substantially equal to the pressure of stage d)" should be understood to mean that the pressure difference of the pressure of stage d) relative to the pressure of stage d) is between 0 and 1 MPa, preferably between 0.005 and 0.3 MPa, and particularly preferably between 0.01 and 0.3 MPa. Preferably, the pressure of stage e) is the pressure of stage d) minus the pressure drop. Operating at least a portion of the separation stage e) at a pressure substantially equal to the operating pressure of stage d) is also beneficial for the recycle of the hydrogen contained in the gaseous effluent.

[0199] The separation stage e) can also be carried out at a pressure below the pressure of stage d).

[0200] The separation stage e) can also include a (first) separation stage which is carried out at a pressure substantially equal to the operating pressure of stage d), followed by at least one other separation stage at a temperature equal to or lower than the temperature of each previous separation stage of stage e) and at a pressure lower than the pressure of each previous separation stage of stage e).

[0201] The separation stage e) can advantageously be carried out by any method known to those skilled in the art, such as, for example, a combination of one or more separators (drums) and / or one or more stripping columns, and a stripping gas, such as a hydrogen-rich feed stream, can optionally be fed to this or these separators (drums) and / or columns. Preferably, stage c) is carried out using a single separator (drum).

[0202] The separation stage e) is preferably carried out in at least one separator drum, which is referred to as a high-pressure or medium-pressure low-temperature separator drum and is also known to those skilled in the art by the name CHPS (cold high-pressure separator). Thus, this stage e) preferably employs a "cold high-pressure" separator, the pressure of which is substantially equal to the operating pressure of stage d).

[0203] The washing / separation stage of stage e) can be carried out at least in part in common or separate washing and separation equipment parts, which are well-known (knockout drums, pumps, heat exchangers, washing towers, etc. that can operate at various pressures and temperatures). The separation stage e) can, for example, include a tower for stripping acidic water from the withdrawn aqueous effluent (also known as an acid water stripper), a tower for washing the acidic gas to purify the hydrogen-rich gas before recycling, and / or a tower for stabilizing the washed liquid effluent to remove dissolved gases.

[0204] The gas effluent obtained at the end of stage e) advantageously contains hydrogen, and preferably contains at least 80% by volume, preferably at least 85% by volume of hydrogen. Advantageously, the gas effluent can be at least partially recycled to the hydrogenation stage a) and / or the hydrotreating stage b) and / or the hydrocracking stage g) (if it exists), and the recycling system can include a purification section.

[0205] The gas effluent can also undergo additional separation to recover at least one hydrogen-rich gas and / or light hydrocarbons, especially ethane, propane, and butane, which can advantageously be sent, individually or as a mixture, to one or more furnaces of the steam cracking stage h), thereby increasing the total olefin yield.

[0206] The aqueous effluent obtained at the end of stage e) advantageously contains ammonium salts and / or hydrogen halides (e.g., hydrochloric acid).

[0207] Depending on the content of halogen compounds in the initial feedstock to be treated, a stream containing amines (such as, for example, monoethanolamine, diethanolamine, and / or mono- and diethanolamine) can be injected upstream of the hydrogenation stage a) and / or the hydrotreating stage c) to ensure that a sufficient amount of ammonium ions binds to the halide ions formed during the hydrogenation and / or hydrotreating processes, thereby making it possible to limit the formation of hydrochloric acid and thus limit the corrosion downstream of the separation section. This injection may be particularly necessary when the feedstock contains many halogen compounds or few nitrogen compounds.

[0208] In an optional embodiment of the present invention, the separation stage e) comprises injecting an aqueous solution into a mixture of the gaseous effluent and at least a portion of the second liquid effluent obtained from stage d), followed by entering a washing / separation section which advantageously comprises a separation phase that allows obtaining at least one aqueous effluent loaded with dissolved ammonium halide salt, a washed liquid hydrocarbon effluent, and a partially washed gaseous effluent. Subsequently, the aqueous effluent loaded with dissolved ammonium halide salt and the washed liquid hydrocarbon effluent can be separated in a settling drum to obtain the hydrocarbon effluent and the aqueous effluent. The partially washed gaseous effluent can be introduced in parallel into a washing tower where it is circulated countercurrently with an aqueous stream preferably of the same type as the aqueous solution injected into the hydrocarbon effluent, so that hydrogen halide (HCl type) contained in the partially washed gaseous effluent can be removed at least partially and preferably completely, and thereby obtain the gaseous effluent (preferably mainly comprising hydrogen) and an acidic aqueous stream. The aqueous effluent obtained from the settling drum can optionally be mixed with the acidic aqueous stream and optionally used as a mixture with the acidic aqueous stream in a water recirculation loop for feeding the aqueous solution upstream of the washing / separation section and / or the aqueous stream in the washing tower to the separation stage e). The water recirculation loop can comprise a water supply and / or a supply of an alkaline solution and / or a bleed so that dissolved salts can be discharged.

[0209] The hydrocarbon effluent from the separation stage e) can alternatively be sent in part or in whole to the fractionation stage f), or directly to the steam cracking stage h), or directly to a fuel storage unit.

[0210] According to another embodiment, at least a portion of the hydrocarbon effluent from the separation stage e) is recycled upstream of stage a) and / or upstream of stage c).

[0211] Mixing the feedstock with the hydrocarbon effluent (thermal recycle) from stage e) upstream of stage a) and / or stage c) allows on the one hand diluting the impurities in the feedstock and on the other hand controlling the temperature of stage a) and / or c) where highly exothermic reactions occur. Mixing the feedstock with a portion (thermal recycle) of the hydrocarbon effluent from the separation stage e) upstream of stage a) also allows indirectly heating the feedstock by simply mixing the "cold" feedstock with the thermal recycle.

[0212] Recycling at least a portion of the hydrocarbon effluent from the separation stage e) also has the advantage that the first gaseous effluent (naphtha fraction) from stage b) bypasses the hydrotreating stage c) and is contained in the hydrocarbon effluent from the separation stage e), which can be hydrotreated via this recycle, especially when at least a portion of the hydrocarbon effluent from the separation stage e) is introduced upstream of the hydrotreating stage c).

[0213] Adjust the amount of the recycled hydrocarbon effluent from separation stage e) such that the weight ratio of the recycle stream from stage e) to the feedstock containing pyrolysis oil (i.e., the feedstock to be treated in the overall process) is less than or equal to 10, preferably less than or equal to 7, and more preferably greater than or equal to 0.001, preferably greater than or equal to 0.01, and preferably greater than or equal to 0.1. Preferably, adjust the amount of the recycled hydrocarbon effluent from separation stage e) such that the weight ratio of the recycle stream to the feedstock containing plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5.

[0214] When a part of the second liquid effluent from stage d) is recycled upstream of stage a) and a part of the hydrocarbon effluent from separation stage e) is recycled upstream of stage a) and / or upstream of stage c), the weight ratio of the total recycle stream (the effluent from stage d) and the effluent from stage e)) to the feedstock containing plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5.

[0215] (f) Fractionation stage (optional)

[0216] The method according to the invention may comprise a stage of fractionating all or part, preferably all, of the hydrocarbon effluent obtained from stage e) to obtain at least a fourth gas effluent, a naphtha fraction and at least one middle distillate oil fraction.

[0217] The term "naphtha fraction" is understood to mean a hydrocarbon fraction containing compounds having a boiling point generally less than or equal to 175 °C, in particular between 80 °C and 175 °C.

[0218] The term "middle distillate oil fraction" is understood to mean a hydrocarbon fraction containing compounds having a boiling point generally greater than 175 °C. The heavy fraction may include middle distillate oils such as diesel fractions and / or kerosene fractions. It may also contain heavier compounds.

[0219] Depending on the purpose or use of the fractions obtained from fractionation stage f), the person skilled in the art will adjust the fractionation points in the stripping and / or distillation operations. For example, it may be necessary to adjust the end point of the naphtha fraction to 150, 175 or 200 °C.

[0220] Stage f) particularly makes it possible to remove gases dissolved in the hydrotreated liquid effluent, such as ammonia, hydrogen sulfide and light hydrocarbons having 1 to 4 carbon atoms.

[0221] Optional fractionation stage f) advantageously operates at a pressure less than or equal to 1.0 MPa absolute, preferably between 0.1 and 1.0 MPa absolute.

[0222] According to one embodiment, stage f) can be operated in a section advantageously comprising at least one stripping column equipped with a reflux circuit comprising a reflux drum. The stripping column is fed with the liquid hydrocarbon effluent obtained from stage e) and a vapor stream. The liquid hydrocarbon effluent obtained from stage e) can optionally be heated before entering the stripping column. Thus, the lightest compounds are entrained at the top of the column and enter the reflux circuit comprising the reflux drum, where a gas / liquid separation takes place. The gas phase comprising light hydrocarbons is withdrawn from the reflux drum as a gas stream. The naphtha fraction is advantageously withdrawn from the reflux drum. The middle distillate fraction is advantageously withdrawn from the bottom of the stripping column.

[0223] According to another embodiment, stage e) can be carried out in a section comprising one or more separation drums. Advantageously, at least one "low pressure low temperature" separation drum will be used to remove the gases dissolved in the liquid hydrocarbon effluent.

[0224] According to other embodiments, the fractionation stage f) can use a stripping column or a separation drum, followed by a distillation column, or only a distillation column.

[0225] The optionally mixed naphtha fraction and middle distillate fraction can be sent in whole or in part to a steam cracking unit, where olefins can be (re)formed at the outlet of the unit to participate in the formation of polymers. Preferably, only a portion of the fractions is sent to the steam cracking unit; at least a portion of the remainder is optionally recycled to at least one stage of the process and / or sent to a fuel storage unit obtained from a conventional petroleum-based feedstock, such as a naphtha storage unit, a diesel storage unit or a kerosene storage unit.

[0226] According to a preferred embodiment, the naphtha fraction is sent in whole or in part to the steam cracking unit, while the middle distillate fraction is sent to the hydrocracking stage g), and / or to a fuel storage unit.

[0227] In a particular embodiment, the optional fractionation stage f) can result in the obtaining of, in addition to the gas stream, a naphtha fraction (generally comprising compounds with a boiling point less than or equal to 175 °C, preferably between 80 and 175 °C) and a middle distillate fraction (generally comprising compounds with a boiling point greater than 175 °C and less than 385 °C), as well as a heavy hydrocarbon fraction (generally comprising compounds with a boiling point greater than or equal to 385 °C). The naphtha fraction can be sent in whole or in part to the steam cracking unit and / or to a naphtha storage unit obtained from a conventional petroleum feedstock; it can also be recycled; the middle distillate fraction can also be sent in whole or in part to the steam cracking unit, or to a diesel storage unit obtained from a conventional petroleum feedstock, or to the hydrocracking stage g) (when it is present), or can also be recycled; the heavy fraction itself can be sent at least in part to the steam cracking unit, or to the hydrocracking stage (when it is present).

[0228] In another specific embodiment, the optional fractionation stage f) can result in obtaining, in addition to the gas stream, a naphtha fraction (usually containing compounds with a boiling point less than or equal to 175 °C, preferably between 80 and 175 °C) and a kerosene fraction (usually containing compounds with a boiling point greater than 175 °C and less than or equal to 280 °C), a diesel fraction (usually containing compounds with a boiling point greater than 280 °C and less than 385 °C), and a heavy hydrocarbon fraction (usually containing compounds with a boiling point greater than or equal to 385 °C). The naphtha fraction, the kerosene fraction, and / or the diesel fraction can be sent in whole or in part to a steam cracking unit, or separately to a naphtha, kerosene, or diesel pool obtained from conventional petroleum feedstocks, or can be recycled. The diesel and / or kerosene fraction can also be sent to the hydrocracking stage g) (when it is present). The heavy fraction itself can be sent at least in part to a steam cracking unit, or to the hydrocracking stage g) (when it is present).

[0229] In another specific embodiment, the naphtha fraction obtained from stage f) is fractionated into a heavy naphtha fraction (usually containing compounds with a boiling point between 80 °C and 175 °C) and a light naphtha fraction (usually containing compounds with a boiling point less than 80 °C), and at least a part of the heavy naphtha fraction is sent to an aromatics complex comprising at least one naphtha reforming stage to produce aromatic compounds. According to this embodiment, at least a part of the light naphtha fraction is sent to the following steam cracking stage h).

[0230] When at least one of these fractions is recycled, it is advantageously recycled upstream of stage a) and / or upstream of stage c).

[0231] The amount of the recycled fraction is adjusted such that the weight ratio of the recycled stream to the feedstock containing pyrolysis oil (i.e., the feedstock to be treated for the entire process) is less than or equal to 10, preferably less than or equal to 7, and more preferably greater than or equal to 0.001, preferably greater than or equal to 0.01, and preferably greater than or equal to 0.1. Preferably, the amount of the recycled fraction is adjusted such that the weight ratio of the recycled stream to the feedstock containing plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5.

[0232] When one or more fractions of the fractionation stage f) and / or a part of the second liquid effluent from stage d) and / or a part of the hydrocarbon effluent from the separation stage e) are recycled, the weight ratio of the total recycled stream to the feedstock containing plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5.

[0233] The gaseous effluent obtained from fractionation stage f) may undergo additional purification and separation to recover at least light hydrocarbons, in particular ethane, propane and butane, which may advantageously be sent, individually or as a mixture, to one or more furnaces of steam cracking stage h) to increase the overall yield of olefins.

[0234] (g) Hydrocracking stage (optional)

[0235] According to one variant, the process of the invention may include a hydrocracking stage g) carried out with at least a portion of the hydrotreated effluent obtained from stage c) after the hydrotreating stage c), or with at least a portion of the middle distillate fraction after the fractionation stage f).

[0236] Advantageously, stage g) uses hydrocracking reactions known to those skilled in the art and more particularly makes it possible to convert heavy compounds, such as compounds having a boiling point greater than 175 °C, into compounds having a boiling point less than or equal to 175 °C, which are included in the hydrocarbon effluent obtained from fractionation stage f). Other reactions may be carried out, such as olefin or aromatic hydrogenation, hydrodemetallization, hydrodesulfurization, hydrodenitrogenation, etc.

[0237] Compounds having a boiling point greater than 175 °C contain more naphthenic compounds, naphtheno-aromatic compounds and aromatic compounds relative to the lighter compounds, thereby resulting in a higher C / H ratio. This high ratio is the cause of coking in steam cracking units and thus requires steam cracking furnaces dedicated to this fraction. When it is desired to minimize the yield of these heavy compounds (middle distillate fraction) and maximize the yield of light compounds (naphtha fraction), these compounds can be at least partially converted into light compounds by hydrocracking - a fraction that is generally advantageous for the steam cracking unit.

[0238] Thus, the process of the invention may include a hydrocracking stage g) carried out in a hydrocracking reaction section using at least one fixed bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed containing at least one hydrocracking catalyst, feeding to the hydrocracking reaction section at least a portion of the hydrotreated effluent obtained from stage c) and / or feeding at least a portion of the middle distillate fraction obtained from stage f) and a third gas stream containing hydrogen, the hydrocracking reaction section being used at an average temperature between 250 and 450 °C, a hydrogen partial pressure between 1.5 and 20.0 MPa absolute pressure and a space velocity between 0.1 and 10.0 h -1 to obtain a first hydrocracking effluent.

[0239] Therefore, the hydrocracking reaction section is advantageously carried out at an average temperature between 250 °C and 480 °C, preferably between 320 °C and 450 °C, at a hydrogen partial pressure between 1.5 and 20.0 MPa absolute, preferably between 3 and 18.0 MPa absolute, and at a space velocity (HSV) between 0.1 and 10.0 h -1 between, preferably between 0.1 and 5.0 h -1 between, preferably between 0.2 and 4 h -1 between. The hydrogen coverage in stage g) is advantageously between 80 and 2000 Sm 3 hydrogen / m 3 between the fresh feedstock fed to stage a) and, preferably, between 200 and 1800 Sm 3 hydrogen / m 3 between the fresh feedstock fed to stage a). The definitions of the average temperature (WABT), HSV and hydrogen coverage correspond to those described above.

[0240] Advantageously, the hydrocracking reaction section is used at a pressure equal to that used in the reaction section of the hydrogenation stage a) or the hydrotreating stage c).

[0241] Advantageously, stage g) is carried out in a hydrocracking reaction section comprising at least one, preferably 1 to 5, fixed-bed reactors having n catalyst beds, where n is an integer greater than or equal to 1, preferably 1 to 10, preferably 2 to 5, and each bed contains at least one, and preferably no more than 10, hydrocracking catalysts. When the reactor contains a plurality of catalyst beds, i.e. at least 2, preferably 2 to 10, preferably 2 to 5, catalyst beds, the catalyst beds are preferably arranged in series in the reactor.

[0242] The effluent from the hydrocracking can be recycled at least in part to the hydrogenation stage a) and / or recycled to the thermal separation stage b) and / or the hydrotreating stage c) and / or the thermal separation stage d) and / or the cold separation stage e) and / or the fractionation stage f).

[0243] The hydrocracking stage can be carried out in one stage (stage g) or two stages (stages g) and g’)). When it is carried out in two stages, the effluent from the first hydrocracking stage g) is separated so that a hydrocarbon fraction (middle distillate fraction) containing compounds with a boiling point greater than 175 °C can be obtained, and this fraction is introduced into the second hydrocracking stage g’), which contains a dedicated second hydrocracking reaction section different from the first hydrocracking reaction section g). This configuration is particularly suitable when only a naphtha fraction needs to be produced.

[0244] The second hydrocracking step g') is carried out in a hydrocracking reaction section using at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, and each catalytic bed contains at least one hydrocracking catalyst. At least a portion of the first hydrocracking effluent obtained from the first hydrocracking stage g) and a gas stream containing hydrogen are fed to the hydrocracking reaction section. The hydrocracking reaction section is operated at an average temperature between 250 and 450 °C, a hydrogen partial pressure between 1.5 and 20.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a second hydrocracking effluent. The preferred operating conditions and catalysts used in the second hydrocracking stage are those described for the first hydrocracking stage. The operating conditions and catalysts used in these two hydrocracking stages can be the same or different.

[0245] The second hydrocracking stage is preferably carried out in a hydrocracking reaction section comprising at least one, preferably 1 to 5, fixed-bed reactors having n catalytic beds, where n is an integer greater than or equal to 1, preferably 1 to 10, more preferably 2 to 5, and each bed contains at least one, preferably no more than 10, hydrocracking catalysts.

[0246] These operating conditions used in the hydrocracking stage generally enable a single-pass conversion of greater than 15 wt%, and more preferably between 20 wt% and 95 wt%, to a product that is converted to compounds having a boiling point of less than or equal to 175 °C, preferably less than 160 °C, and more preferably less than 150 °C, with at least 80 vol%. When the process is carried out in two hydrocracking stages, the single-pass conversion in the second stage is kept moderate to maximize the selectivity for compounds in the naphtha fraction (boiling point less than or equal to 175 °C, especially between 80 and less than or equal to 175 °C). The single-pass conversion is limited by using a high recycle ratio in the loop of the second hydrocracking stage. The recycle ratio is defined as the ratio of the feed flow rate of stage g') to the feed flow rate of stage a); preferably, this ratio is between 0.2 and 4, more preferably between 0.5 and 2.5.

[0247] The hydrocracking effluent from the second hydrocracking stage g') can be at least partially recycled to the hydrogenation stage a) and / or recycled to the thermal separation stage b) and / or the hydrotreating stage c) and / or the thermal separation stage d) and / or the cold separation stage d) and / or the fractionation stage e).

[0248] The hydrocracking stage thus does not necessarily convert all hydrocarbon compounds of the middle distillate fraction into hydrocarbon compounds with a boiling point less than or equal to 175 °C (naphtha fraction). After the fractionation stage f), there may thus remain a more or less significant proportion of compounds with a boiling point greater than 175 °C. At least a portion of this unconverted fraction can be introduced into a second hydrocracking stage g’). Another portion can be discharged. Depending on the operating conditions of the process, the bleed can be a fraction containing compounds with a boiling point greater than 175 °C between 0 wt% and 10 wt%, and preferably between 0.5 wt% and 5 wt% relative to the input feedstock.

[0249] According to the invention, the hydrocracking stage is carried out in the presence of at least one hydrocracking catalyst.

[0250] The hydrocracking catalyst used in the hydrocracking stage is a conventional bifunctional hydrocracking catalyst known to those skilled in the art, which combines an acidic function with a hydrodehydrogenation function and optionally at least one binder matrix. The acidic function is contributed by a high surface area (usually 150 to 800 m 2 / g) support, such as halogenated (especially chlorinated or fluorinated) alumina, a combination of aluminum and boron oxides, amorphous silica-alumina, and zeolites. The hydrodehydrogenation function is contributed by at least one metal from Group VIB of the periodic table and / or at least one metal from Group VIII.

[0251] Preferably, the hydrocracking catalyst comprises a hydrodehydrogenation function containing at least one Group VIII metal selected from iron, cobalt, nickel, ruthenium, rhodium, palladium, and platinum, and preferably selected from cobalt and nickel. Preferably, the catalyst also comprises at least one Group VIB metal selected from chromium, molybdenum, and tungsten (alone or as a mixture), and preferably selected from molybdenum and tungsten. Hydrodehydrogenation functions of the NiMo, NiMoW, or NiW type are preferred.

[0252] Preferably, the content of the Group VIII metal in the hydrocracking catalyst is advantageously between 0.5 wt% and 15 wt%, and preferably between 1 wt% and 10 wt%, the percentage being expressed as the weight percentage of the oxide relative to the total weight of the catalyst. When the metal is cobalt or nickel, the metal content is expressed as CoO and NiO, respectively.

[0253] Preferably, the content of the Group VIB metal in the hydrocracking catalyst is advantageously between 5 wt% and 35 wt%, and preferably between 10 wt% and 30 wt%, the percentage being expressed as the weight percentage of the oxide relative to the total weight of the catalyst. When the metal is molybdenum or tungsten, the metal content is expressed as MoO3 and WO3, respectively.

[0254] The hydrocracking catalyst may also optionally contain at least one promoter element deposited on the catalyst and selected from phosphorus, boron, and silicon, optionally at least one element of Group VIIA (chlorine and fluorine are preferred), optionally at least one element of Group VIIB (manganese is preferred), and optionally at least one element of Group VB (niobium is preferred).

[0255] Preferably, the hydrocracking catalyst comprises at least one amorphous or weakly crystalline porous inorganic matrix of an oxide type selected from alumina, silica, silica-alumina, aluminate, alumina-boron oxide, magnesia, silica-magnesia, zirconia, titania, or clay (alone or as a mixture), and preferably alumina or silica-alumina (alone or as a mixture).

[0256] Preferably, the silica-alumina contains more than 50 wt% of alumina, preferably more than 60 wt% of alumina.

[0257] Preferably, the hydrocracking catalyst also optionally contains a zeolite selected from Y zeolites, preferably USY zeolites, alone or in combination with other zeolites selected from beta zeolite, ZSM-12, IZM-2, ZSM-22, ZSM-23, SAPO-11, ZSM-48, or ZBM-30 zeolites (alone or as a mixture). Preferably, the zeolite is a single USY zeolite.

[0258] In the case where the catalyst contains a zeolite, the zeolite content in the hydrocracking catalyst is advantageously between 0.1 wt% and 80 wt%, preferably between 3 wt% and 70 wt%, the percentages being expressed as percentages of zeolite relative to the total weight of the catalyst.

[0259] The preferred catalyst comprises at least one Group VIB metal and optionally at least one Group VIII non-noble metal, at least one promoter element, and preferably phosphorus, at least one Y zeolite, and at least one alumina binder, and preferably consists of them.

[0260] A more preferred catalyst comprises nickel, molybdenum, phosphorus, USY zeolite, and optionally beta zeolite and alumina, and preferably consists of them.

[0261] Another preferred catalyst comprises nickel, tungsten, alumina, and silica-alumina, and preferably consists of them.

[0262] Another preferred catalyst comprises nickel, tungsten, USY zeolite, alumina, and silica-alumina, and preferably consists of them.

[0263] The hydrocracking catalyst is in the form of, for example, extrudates.

[0264] In one variant, the hydrocracking catalyst used in the second hydrocracking stage comprises a hydrodehydrogenation functionality containing at least one Group VIII noble metal selected from palladium and platinum (alone or as a mixture). The content of the Group VIII noble metal is advantageously between 0.01% by weight and 5% by weight, and preferably between 0.05% by weight and 3% by weight, the percentages being expressed as the weight percentage of the oxide (PtO or PdO) relative to the total weight of the catalyst.

[0265] According to another aspect of the invention, the hydrocracking catalyst additionally comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur. Such catalysts are generally denoted by the term "additivated catalyst". Generally, the organic compounds are selected from compounds containing one or more chemical functionalities selected from carboxyl, alcohol, thiol, thioether, sulfone, sulfoxide, ether, aldehyde, ketone, ester, carbonate, amine, nitrile, imide, oxime, urea and amide functionalities, or compounds containing a furan ring or sugar.

[0266] The preparation of the catalysts for the hydrogenation, hydrotreating and hydrocracking stages is known and generally comprises the stages of impregnating a support with a Group VIII metal and a Group VIB metal (when present) and optionally phosphorus and / or boron, then drying, and then optionally calcining. In the case of additivated catalysts, this preparation is generally carried out by simple drying without calcining after introducing the organic compound. The term "calcining" is understood herein to mean a heat treatment at a temperature greater than or equal to 200 °C in a gas containing air or oxygen. Before their use in the stages of the process, the catalysts are generally subjected to sulfidation to form the active mass. The catalyst of stage a) can also be a catalyst used in its reduced form, so a reduction stage is involved in its preparation.

[0267] The gas stream containing hydrogen fed to the hydrogenation, hydrotreating and hydrocracking reaction sections can consist of a hydrogen make-up, and / or can consist of recycled hydrogen obtained in particular from the separation stage e) or from the fractionation stage f). Preferably, an additional gas stream containing hydrogen is advantageously introduced at the inlet of each reactor (in particular operating in series) and / or at the inlet of each catalytic bed starting from the second catalytic bed of the reaction section. These additional gas streams are also called cooling streams. They make it possible to control the temperature in the reactors where the reactions taking place are generally highly exothermic.

[0268] The hydrotreated effluent or the hydrocarbon fraction thus obtained by treating the pyrolysis oil by the method according to the invention exhibits a composition compatible with the feedstock specifications at the inlet of a steam cracking unit. In particular, the composition of the hydrotreated effluent or the hydrocarbon fraction is preferably such that:

[0269] - The total content of metallic elements is less than or equal to 10.0 weight ppm, preferably less than or equal to 2.0 weight ppm, preferably less than or equal to 1.0 weight ppm, and preferably less than or equal to 0.8 weight ppm, where:

[0270] The content of silicon (Si) element is less than or equal to 5.0 weight ppm, preferably less than or equal to 1 weight ppm, and preferably less than or equal to 0.6 weight ppm, and

[0271] The content of iron (Fe) element is less than or equal to 200 weight ppb,

[0272] - The sulfur content is less than or equal to 500 weight ppm, preferably less than or equal to 200 weight ppm,

[0273] - The nitrogen content is less than or equal to 100 weight ppm, preferably less than or equal to 50 weight ppm, and preferably less than or equal to 5 weight ppm,

[0274] - The total content of chlorine element is less than or equal to 5.0 weight ppm, preferably less than 1.0 weight ppm,

[0275] - The content of olefin compounds (monoolefins and diolefins) is less than or equal to 5.0 weight %, preferably less than or equal to 2.0 weight %, and preferably less than or equal to 0.1 weight %,

[0276] - The mercury content is less than or equal to 5 weight ppb, preferably less than 3 weight ppb.

[0277] The content is given as relative weight concentration, weight percentage (%), weight parts per million (ppm) or weight parts per billion (ppb) relative to the total weight of the stream under consideration.

[0278] The process according to the invention thus makes it possible to treat pyrolysis oil to obtain at least one effluent that can be injected in whole or in part into at least one steam cracking unit.

[0279] Heavy metal adsorption stage (optional)

[0280] At least one gas effluent and / or any liquid effluent obtained from separation stages b), b'), d) and e) and / or any gas effluent and / or any liquid effluent obtained from fractionation stage f) can be subjected to an optional heavy metal adsorption stage.

[0281] The gas effluent is in particular the first gas effluent obtained from stage b) and / or the gas effluent obtained from stage b') and / or the second gas effluent obtained from stage d) and / or the third gas effluent obtained from stage e) and / or the fourth gas effluent obtained from stage f).

[0282] Liquid effluents, in particular the first liquid effluent obtained from stage b), and / or the hydrocarbon effluent obtained from stage b'), and / or the second liquid effluent obtained from stage d), the hydrocarbon effluent obtained from stage e) and / or at least one fraction obtained from stage f).

[0283] An optional adsorption stage makes it possible to remove or reduce the amount of metal impurities that may be present in said gas and liquid effluents, in particular heavy metals such as arsenic, zinc, lead and especially mercury. Metal impurities can be present in the feedstock and / or formed during the process of the method stages. Their removal or reduction may be particularly necessary when at least a portion of said gas and liquid effluents is intended to be sent directly or, after undergoing one or more optional additional stages (such as fractionation stage f)), to a steam cracking stage. Regulations regarding metal impurities, especially mercury, at the inlet of the steam cracking stage may require such a stage.

[0284] Thus, the optional adsorption stage is advantageously carried out, especially when at least one of these effluents or feedstocks contains respectively more than 20 weight ppb, especially more than 15 weight ppm of heavy metals (As, Zn, Pb, Hg, etc.), and especially when at least one of these effluents or feedstocks contains respectively more than 10 weight ppm of mercury, more particularly more than 15 weight ppm of mercury.

[0285] The optional adsorption stage is advantageously carried out at a temperature between 20 and 150 °C, preferably between 40 and 100 °C and at an absolute pressure between 0.15 and 10.0 MPa, preferably between 0.2 and 1.0 MPa.

[0286] The optional adsorption stage can be carried out with any adsorbent known to those skilled in the art that makes it possible to reduce the amount of such contaminants.

[0287] According to a variant, the optional adsorption stage is carried out in an adsorption section that operates in the presence of at least one adsorbent comprising a porous support and at least one active phase based on sulfur in elemental form or sulfur in metal sulfide form.

[0288] The porous support can be chosen indifferently from the alumina, silica-alumina, silica, zeolite or activated carbon families. Advantageously, the porous support is based on alumina. The specific surface area of the support is generally between 150 and 600 m 2 / g, preferably between 200 and 400 m 2 / g, even more preferably between 150 and 320 m 2 / g. The specific surface area of the adsorbent is the surface area measured by the BET method as described above.

[0289] The active phase is based on sulfur in elemental form or sulfur in the form of metal sulfides, in particular sulfides of metals selected from copper, molybdenum, tungsten, iron, nickel or cobalt.

[0290] Advantageously, based on the total weight of the adsorbent, the active phase of the adsorbent contains 1% to 70% by weight of sulfur, preferably 2% to 25%, and very preferably 3% to 20%. Based on the total weight of the adsorbent, the weight proportion of the metal is generally between 1% and 60%, preferably between 2% and 40%, preferably between 5% and 30%, and very preferably between 5% and 20%.

[0291] The residence time in the adsorption section is generally between 1 and 180 minutes.

[0292] The adsorption section comprises at least one adsorption tower, preferably at least two adsorption towers, preferably 2 to 4 adsorption towers, which contain the adsorbent. When the adsorption section comprises two adsorption towers, one operating mode can be "swing" operation according to the specific terminology, where one tower is on-line, i.e., in operation, while the other tower is in standby. Another operating mode is to operate at least two towers in series in a replaceable mode.

[0293] (h) Steam cracking stage (optional)

[0294] The hydrocarbon effluent obtained from the cold separation stage e), or at least one liquid hydrocarbon fraction obtained from the optional stage f) can be sent in whole or in part to the steam cracking stage h).

[0295] Advantageously, the gas effluent obtained from the separation stages b), b'), d) and / or e) and / or the fractionation stage f) and containing ethane, propane and butane can also be sent in whole or in part to the steam cracking stage h).

[0296] The steam cracking stage h) is advantageously carried out in at least one pyrolysis furnace at a temperature between 700 and 900 °C, preferably between 750 and 850 °C, and at a relative pressure of 0.05 to 0.3 MPa. The residence time of the hydrocarbon compounds is generally less than or equal to 1.0 second (expressed as s), preferably between 0.1 and 0.5 seconds. Advantageously, steam is introduced upstream of the optional steam cracking stage h) and after separation (or fractionation). The amount of water introduced (advantageously in the form of steam) is advantageously 0.3 to 3.0 kg of water / kg of hydrocarbon compounds at the inlet of stage h). Preferably, the optional stage h) is carried out in a plurality of pyrolysis furnaces in parallel to adapt the operating conditions to the feeds to stage h), in particular the various streams obtained from stage f), and also to manage the decoking time of the tubes. The furnace comprises one or more tubes arranged in parallel. The furnace can also refer to a set of furnaces operating in parallel. For example, the furnace can be dedicated to the cracking of middle distillate fractions.

[0297] The effluents from various steam crackers are generally recombined before separation to form the effluent. It is to be understood that the steam cracking stage h) includes steam crackers but also sub-stages related to steam cracking known to those skilled in the art. These sub-stages may in particular include heat exchangers, columns and catalytic reactors as well as operations for recycling to the furnace. The column generally makes it possible to fractionate the effluent to at least recover a light fraction containing hydrogen and compounds having 2 to 5 carbon atoms, a fraction containing pyrolysis gasoline, and optionally, heavier fractions. The column makes it possible to separate the various components of this fractionated light fraction to at least recover an ethylene-rich fraction (C2 fraction) and a propylene-rich fraction (C3 fraction) and optionally a butene-rich fraction (C4 fraction). The catalytic reactor in particular makes it possible to carry out the hydrogenation of the C2, C3 and even C4 fractions as well as pyrolysis gasoline. The saturated compounds, in particular those having 2 to 4 carbon atoms, are advantageously recycled to the steam cracker to increase the total yield of olefins.

[0298] This steam cracking stage h) makes it possible to obtain at least one effluent containing a satisfactory content of olefins having 2, 3 and / or 4 carbon atoms (i.e., C2, C3 and / or C4 olefins), in particular greater than or equal to 30% by weight, in particular greater than or equal to 40% by weight, or even greater than or equal to 50% by weight of the total olefins having 2, 3 and 4 carbon atoms, relative to the weight of the steam cracking effluent under consideration. The C2, C3 and C4 olefins can then be advantageously used as polyolefin monomers.

[0299] The analytical methods used

[0300] The analytical methods and / or standards for determining the characteristics of the various streams, in particular the characteristics of the feedstock to be treated and the effluent, are known to those skilled in the art. They are listed in particular below for information. Other methods considered equivalent can also be used, in particular equivalent IP, EN or ISO methods:

[0301] Table 1

[0302]

[0303] (1) The MAV method is described in the article: C. López-García et al., Near Infrared Monitoring of Low Conjugated Diolefins Content in Hydrotreated FCC Gasoline Streams, Oil & Gas Science and Technology – Rev. IFP, Vol. 62 (2007), No. 1, pp. 57-68.

[0304] List of figures

[0305] Figures 1 to 4 The details of the components mentioned make it possible to better understand the present invention, which is not limited to Figures 1 to 4 the specific embodiments shown. The various embodiments shown can be used alone or in combination with each other, and there is no limitation on such combination. In the drawings, the same reference numerals denote the same or similar elements.

[0306] Figure 1 A diagram representing a specific embodiment of the method of the present invention, which comprises:

[0307] - Stage a), hydrogenation of a hydrocarbon feedstock 1 obtained from plastic pyrolysis, preferably as a mixture with at least a part of a recycled second liquid effluent 11a obtained from stage d), and in the presence of a hydrogen-rich gas 2 and optionally an amine contributed by a stream 3 and optionally a sulfur compound contributed by a stream 4, which is carried out in at least one fixed-bed reactor containing at least one hydrogenation catalyst to obtain a hydrogenated effluent 5;

[0308] - Stage b), separation of the hydrogenated effluent 5 carried out at high pressure and high temperature (HHPS) to obtain at least a first gas effluent 6 and a first liquid effluent 7,

[0309] - Stage c), hydrotreating of the first liquid effluent 7 obtained from stage b) in at least one fixed-bed reactor containing at least one hydrotreating catalyst in the presence of hydrogen 8 to obtain a hydrotreated effluent 9;

[0310] - Stage d), separation of the hydrotreated effluent 9 at high pressure and high temperature (HHPS) to obtain at least a second gas effluent 10 and a second liquid effluent 11, at least a part 11a of the second liquid effluent 11 being preferably recycled upstream of stage a);

[0311] - Separation stage e), which is carried out at high pressure and low temperature (CHPS) and fed with the first gas effluent 6 and the second gas effluent 10 and at least a part of the second liquid effluent 11 obtained from stage d) and an aqueous solution 12, such that it is possible to obtain at least a third gas effluent 13 containing hydrogen, an aqueous effluent 14 containing dissolved salts and a hydrocarbon effluent 15;

[0312] - Optionally, stage f), fractionation of the hydrocarbon effluent 15 such that it is possible to obtain at least a fourth gas effluent 16, a naphtha fraction 17 (usually containing compounds with a boiling point less than or equal to 175 °C) and an intermediate distillate fraction 18 (usually containing compounds with a boiling point greater than 175 °C).

[0313] At the end of stage f), a portion of the naphtha fraction 17 can be sent to a steam cracking process (not shown). Another portion of the naphtha fraction 17 can be fed to the hydrogenation stage a) and / or the hydrotreating stage c) (not shown).

[0314] Figure 2 Figure representing another specific embodiment of the inventive method based on Figure 1 This figure particularly includes a hydrocracking stage g) after the fractionation stage f), where at least a portion of the middle distillate fraction 18 obtained from stage f) is fed to this hydrocracking stage g), which is carried out in at least one fixed bed reactor containing at least one hydrocracking catalyst and fed with hydrogen 19. The hydrocracking effluent 20 is recycled upstream of the separation stage e).

[0315] Figure 3 Figure representing another specific embodiment of the inventive method based on Figure 1 This embodiment particularly includes a hydrocracking stage g) directly after the hydrotreating stage c), where at least a portion of the hydrotreated effluent 9 obtained from stage c) is fed to this hydrocracking stage g), which is carried out in at least one fixed bed reactor containing at least one hydrocracking catalyst and fed with hydrogen 19. Feeding the hydrotreated effluent from the hydrotreating stage c) directly into the hydrocracking stage has the advantage of not requiring it to be repressurized, thus saving energy. Subsequently, the hydrocracking effluent 20 is sent to the thermal separation stage d).

[0316] Figure 4 Figure representing another specific embodiment of the inventive method based on Figure 1 This embodiment particularly includes a cold separation / washing stage b') after the thermal separation stage b) and before the hydrotreating stage c). Wherein, at least a portion of the first liquid effluent 7 from stage b) is fed to this stage b'). Stage b' is carried out at high pressure and low temperature (CHPS) and is fed with the first liquid effluent 7 and the aqueous solution 21. Thereby, it is possible to obtain at least a gas effluent 22 containing hydrogen, an aqueous effluent 23 containing dissolved salts, and a hydrocarbon effluent 24 introduced into the hydrotreating stage c). In this way, the halide in the form of hydrogen halide dissolved in the first liquid effluent 7 is released into the gas effluent 22 during the cold separation stage b').

[0317] Instead of injecting the amine stream 3 into the inlet of the hydrogenation stage a), it is possible to inject it at the inlet of the hydrotreating stage c), at the inlet of the separation stage e), at the inlet of the hydrocracking stage g) (when it exists), or not inject it, depending on the characteristics of the feedstock.

[0318] In Figures 1 to 4Only the main stages and the main material flows are shown in order to better understand the present invention. It is clearly understood that all the equipment parts required for operation are present (drums, pumps, exchangers, furnaces, towers, etc.), even if they are not shown. It is also understood that the hydrogen-rich gas material flow (makeup or recycle) as described above can be injected at the inlet of each reactor or catalytic bed or between two reactors or two catalytic beds. Means for hydrogen purification and recycle known to those skilled in the art can also be used. Example

[0319] Example 1 is an example not according to the present invention in which there is no first thermal separation stage b) between the hydrogenation stage a) and the hydrotreating stage c).

[0320] Example 2 is an example according to the present invention in which a first thermal separation stage b) is carried out between the hydrogenation stage a) and the hydrotreating stage c).

[0321] The feedstock 1 treated in this process is a plastic pyrolysis oil exhibiting the properties shown in Table 2 (i.e., containing 100 wt% of said plastic pyrolysis oil).

[0322] Table 2: Properties of the feedstock

[0323]

[0324]

[0325] Subject the feedstock 1 to a hydrogenation stage a) carried out in a fixed-bed reactor and under the operating conditions shown in Table 4 in the presence of hydrogen 2 and a NiMo-alumina type hydrogenation catalyst to obtain a hydrogenated effluent 5.

[0326] At the end of the hydrogenation stage a), the observed conversion rate (= (initial concentration - final concentration) / initial concentration) is shown in Table 3.

[0327] Table 3: Conversion rates of entities during the hydrogenation stage a)

[0328] Conversion rate of chlorine ％ >90 Conversion rate of diene ％ >70 Conversion rate of olefin ％ >70 Silicon retention rate ％ >85

[0329] According to Example 1 (not according to the present invention), the hydrogenated effluent 5 obtained from the hydrogenation stage a) is directly subjected, without separation, to the hydrotreating stage c) carried out in a fixed bed and in the presence of hydrogen 8 and a hydrotreating catalyst of the NiMo-alumina type under the conditions presented in Table 4. The hydrotreated effluent is then subjected to a separation stage d) at a pressure substantially equal to that of stage c) and at a temperature controlled at 300 °C so that a gas effluent and a liquid effluent can be obtained. Subsequently, the gas effluent and the liquid effluent are subjected to a cold separation stage e): a water stream is injected into the mixture; the final mixture reaches a temperature of 40 °C in a cold HP drum, which operates at a pressure substantially equal to that of stage d), and a hydrogen-rich gas fraction, an aqueous fraction, and a washed hydrocarbon effluent are obtained at its outlet.

[0330] The hydrotreated effluent is then introduced into a so-called low-pressure low-temperature separation drum to remove dissolved gases. The properties of the hydrotreated effluent free of dissolved gases are presented in Table 4 and are consistent with the feedstock specifications of a steam cracking unit. All or part of the hydrotreated effluent obtained can subsequently be upgraded in a steam cracking stage to form olefins, which can be polymerized to form recycled plastics.

[0331] According to Example 2 (according to the present invention), the hydrogenated effluent 5 from the hydrogenation stage a) is subjected to a (thermal) separation stage b) at a pressure substantially the same as that of stage a) and at a temperature controlled at 300 °C so that a first gas effluent 6 and a first liquid effluent 7 can be obtained.

[0332] The first liquid effluent 7 is then introduced into the hydrotreating stage c), which is carried out in a fixed bed and in the presence of hydrogen 8 and a hydrotreating catalyst of the NiMo-alumina type under the conditions presented in Table 4 to obtain a hydrotreated effluent 9.

[0333] The hydrotreated effluent 9 is then introduced into a second thermal separation stage d) at a pressure substantially equal to that of stage c) and at a temperature controlled at 300 °C so that a second gas effluent 10 and a second liquid effluent 11 can be obtained.

[0334] The first gas effluent 6, the second gas effluent 10, and the second liquid effluent 11 are then subjected to a cold separation stage e): a water stream 12 is injected into the mixture, the final mixture reaches a temperature of 40 °C in a cold HP drum, which operates at a pressure substantially equal to that of stage d), and a hydrogen-rich gas fraction 13, an aqueous fraction 14, and a washed hydrocarbon effluent 15 are obtained at its outlet.

[0335] Subsequently, the hydrotreated effluent 15 is introduced into a so-called low-pressure low-temperature separator drum to remove dissolved gases. The properties of the hydrotreated effluent without dissolved gases are presented in Table 4, and these properties are consistent with the feedstock specifications of the steam cracking unit. Subsequently, all or part of the hydrotreated effluent can be upgraded in the steam cracking stage to form olefins, which can be polymerized to form recycled plastics.

[0336] Table 4: Operating conditions of stages a), b), c) and d) and properties of the hydrotreated effluent

[0337]

[0338]

[0339] It has been observed that, in order to obtain a hydrotreated effluent that meets the feedstock specifications of the steam cracking unit, the average temperature in hydrotreating stage c) according to Example 2 is 8 °C lower than the average temperature according to Example 1.

[0340] When the catalyst deactivates during the course of its catalytic cycle, the average temperature can be increased to compensate for the catalytic deactivation. It is possible to increase the temperature up to the temperature at which the catalyst must be replaced (the temperature at the end of the cycle).

[0341] When the temperature difference between the start and the end of the catalytic cycle is set at 60 °C, the catalyst deactivation rate and the catalytic cycle time can be deduced.

[0342] Table 5 shows the influence of having or not having the thermal separation stage b) on the cycle time of the hydrotreating stage c). It has been observed that the cycle time according to Example 1 is 12 months, while the cycle time according to Example 2 is 20 months.

[0343] Table 5: Catalytic cycle time of the hydrotreating stage with ΔT(SOR*-EOR**)=60 °C

[0344]

[0345] (*)SOR (start of cycle or start of operation)

[0346] (**)EOR (end of cycle or end of operation)

[0347] According to Example 2, the fact that the gas effluent from stage b) bypasses the hydrotreating stage c) allows the reactor size of the hydrotreating stage c) to be reduced, and thus savings in terms of catalyst and equipment requirements are achieved. This is confirmed in Table 6.

[0348] Table 6: Volume of catalyst required for the hydrotreating stage c)

[0349]

Claims

1. A method for treating a feedstock, the feedstock comprising pyrolysis oil from plastics and / or tyres containing halogen compounds, the method comprising: a) The hydrogenation stage carried out in the hydrogenation reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed containing at least one hydrogenation catalyst, feeding to the hydrogenation reaction section at least the feedstock, optionally as a mixture with at least a portion of the second liquid effluent obtained from the separation stage d), and a first gas stream containing hydrogen, the hydrogenation reaction section being used at an average temperature between 140 and 400 °C, a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a hydrogenated effluent, b) a separation stage, to which is fed the hydrogenation effluent obtained from stage a), the stage operating at a temperature above the precipitation temperature of ammonium halide and at a pressure substantially equal to the pressure of stage a), to obtain at least a first gas effluent and a first liquid effluent, c) The hydrotreating stage carried out in the hydrotreating reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed containing at least one hydrotreating catalyst, feeding to the hydrotreating reaction section at least the first liquid effluent obtained from stage b) and a second gas stream containing hydrogen, the hydrotreating reaction section being used at an average temperature between 250 and 430 °C, a hydrogen partial pressure between 1.0 and 10.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a hydrotreated effluent, d) a separation stage, to which is fed the hydrotreated effluent obtained from stage c), the stage operating at a temperature above the precipitation temperature of ammonium halide and at a pressure substantially equal to the pressure of stage c), to obtain at least a second gas effluent and a second liquid effluent, with a portion of the second liquid effluent optionally recycled upstream of stage a), e) a separation stage, to which are fed the first and second gas effluents and at least a portion of the second liquid effluent obtained from stage d) and an aqueous solution, the stage operating at a temperature below the precipitation temperature of ammonium halide and at a pressure substantially equal to or lower than the pressure of stage d) to obtain at least a third gas effluent, an aqueous effluent and a hydrocarbon effluent, f) Optionally, a fractionation stage of all or part of the hydrocarbon effluent obtained from stage e) to obtain at least a fourth gas effluent, a naphtha fraction and at least one middle distillate fraction, g) Optionally, a hydrocracking stage carried out in the hydrocracking reaction section, which uses at least one fixed-bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalytic bed containing at least one hydrocracking catalyst, feeding to the hydrocracking reaction section at least a portion of the hydrotreated effluent obtained from stage c) and / or at least a portion of the middle distillate fraction obtained from stage f) and a third gas stream containing hydrogen, the hydrocracking reaction section being used at an average temperature between 250 and 450 °C, a hydrogen partial pressure between 1.5 and 20.0 MPa absolute pressure, and a space velocity between 0.1 and 10.0 h -1 to obtain a first hydrocracking effluent.

2. The method according to the preceding claim, which includes fractionation stage f).

3. The method according to any one of the preceding claims, which includes hydrocracking stage g).

4. The method according to any one of the preceding claims, which includes a separation stage b') carried out between stages b) and c), to which is fed at least a portion of the first liquid effluent obtained from step b) and an aqueous solution, the stage operating at a temperature below the precipitation temperature of ammonium halide and at a pressure substantially equal to or lower than the pressure of stage b), to obtain at least one gas effluent, an aqueous effluent and a hydrocarbon effluent.

5. The method according to the preceding claim, wherein stage b) or stage d) operates at a temperature between 200 and 450 °C, and stage e) or stage b') operates at a temperature greater than or equal to 20 °C and less than 200 °C.

6. The method according to any one of the preceding claims, wherein at least a portion of the second liquid effluent obtained in stage d) is recycled upstream of stage a).

7. The method according to any one of the preceding claims, wherein at least a portion of the hydrocarbon effluent obtained from stage e) is recycled upstream of stage a) and / or upstream of stage c).

8. The method according to any one of the preceding claims, wherein in stage a), the hydrogen coverage rate is between 250 and 800 Sm 3 hydrogen / m 3 feedstock (Sm 3 / m 3 ).

9. The method according to any one of the preceding claims, which includes at least one stage a0) for pretreating a fraction of the pyrolysis oil from plastics and / or tyres, the pretreatment stage being carried out upstream of stage a) and including an adsorption stage and / or a filtration stage and / or a centrifugation stage and / or a sedimentation stage and / or an electrostatic separation stage and / or a stage of washing with an aqueous solution and / or a gas stripping stage.

10. The method according to any one of the preceding claims, wherein at least one, in whole or in part, of the hydrocarbon effluent obtained from separation stage e) or the fraction obtained from stage f) is sent to a steam cracking stage h) carried out in at least one pyrolysis furnace at a temperature between 700 and 900 °C and at a relative pressure between 0.05 and 0.3 MPa.

11. The method according to any one of the preceding claims, wherein the reaction section of stage a) uses at least two reactors operating in a replaceable mode.

12. The method according to any one of the preceding claims, wherein a stream containing an amine and / or a sulfur compound is injected upstream of stage a).

13. The method according to any one of the preceding claims, wherein the hydrogenation catalyst comprises a support selected from alumina, silica, silica-alumina, magnesia, clay, and mixtures thereof, and a hydrodehydrogenation functionality containing at least one Group VIII element and at least one Group VIB element or a hydrodehydrogenation functionality containing at least one Group VIII element.

14. The method according to any one of the preceding claims, wherein the hydrotreating catalyst comprises a support selected from alumina, silica, silica-alumina, magnesia, clay, and mixtures thereof, and a hydrodehydrogenation functionality containing at least one Group VIII element and / or at least one Group VIB element.

15. The method according to any one of the preceding claims, additionally comprising a second hydrocracking stage g') carried out in the hydrocracking reaction section, which uses at least one fixed bed reactor having n catalytic beds, where n is an integer greater than or equal to 1, each catalyst comprising at least one hydrocracking catalyst, feeding to the hydrocracking reaction section at least a portion of the first hydrocracking effluent obtained from the first hydrocracking stage g) and a gas stream comprising hydrogen, the hydrocracking reaction section being used at a temperature between 250 and 450 °C, a hydrogen partial pressure between 1.5 and 20.0 MPa absolute pressure and a space velocity between 0.1 and 10.0 h -1 to obtain a second hydrocracking effluent.

16. The method according to any one of the preceding claims, wherein the hydrocracking catalyst comprises a support selected from halogenated alumina, a combination of boron and aluminum oxides, amorphous silica-alumina, and zeolite, and a hydrodehydrogenation functionality, the hydrodehydrogenation functionality comprising at least one Group VIB metal selected from chromium, molybdenum, and tungsten, alone or as a mixture, and / or at least one Group VIII metal selected from iron, cobalt, nickel, ruthenium, rhodium, palladium, and platinum.

17. A product obtained by the method according to any one of claims 1 to 16.

18. The product according to claim 17, which comprises, relative to the total weight of the product: - a total metal element content of less than or equal to 10.0 weight ppm, - an iron element content of less than or equal to 200 weight ppb, - a silicon element content of less than or equal to 5.0 weight ppm, - a sulfur content of less than or equal to 500 weight ppm, - a nitrogen content of less than or equal to 100 weight ppm, - a chlorine element content of less than or equal to 10 weight ppm, a mercury content of less than or equal to 5 weight ppb.

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