Methods for treating tire pyrolysis oil

Abstract

This invention relates to a method for processing a feedstock comprising at least 5% by weight of tire pyrolysis oil relative to the total weight of the feedstock, the tire pyrolysis oil having a sulfur content of 2,000 to 10,000 ppm by weight relative to the total weight of the oil, comprising: a) an optional selective hydrogenation step, b) a hydrogenation step in the presence of hydrogen and a catalyst to obtain a hydrogenated effluent, c) a hydrogenation treatment step of the hydrogenated effluent in the presence of hydrogen and a catalyst to obtain a hydrogenated treated effluent, d) a separation step performed between and / or after step b) and step c), the step feed comprising the hydrogenated effluent obtained from step b) and / or the hydrogenated treated effluent obtained from step c), and an aqueous solution to obtain at least one gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

Description

Technical Field

[0001] This invention falls within the field of tire recycling. More specifically, the invention relates to a method for processing a feedstock comprising pyrolysis oil, wherein at least 5% by weight of the pyrolysis oil is derived from tires. Tire pyrolysis oil typically contains a relatively high sulfur content. This method is able to at least partially remove impurities that may be relatively abundant in the feedstock, while reducing the consumption of sulfiding agents required to maintain the catalyst in a sulfidated state during the catalytic step of the method by using pyrolysis oil with a high sulfur compound loading. Background Technology

[0002] Plastics or recycled tires or solid recycled fuels (SRF) obtained from collection and sorting channels can be upgraded through a pyrolysis step. Pyrolysis involves heating feedstocks containing plastics and / or recycled tires and / or SRF under anaerobic conditions, with or without a catalyst, to form three main products: a hydrocarbon phase that is liquid at ambient temperature (also known as pyrolysis oil), a light gas phase (also known as pyrolysis gas), and a solid residue (also known as coke).

[0003] The liquid hydrocarbon phase, namely pyrolysis oil, can be upgraded in units used for storing gasoline, kerosene or diesel fuel.

[0004] Another way to upgrade pyrolysis oil is to use it as feedstock for steam cracking units in order to (re)produce olefins, which are constituent monomers of certain polymers.

[0005] However, pyrolysis oils typically contain high levels of impurities that are incompatible with fuel storage units, steam cracking units, or units downstream of steam cracking units (particularly polymerization and selective hydrogenation processes). These impurities can cause operational problems, particularly corrosion (especially due to the presence of chlorine), coking or catalytic deactivation, or compatibility issues with the target polymer. The presence of dienes can also lead to instability issues in pyrolysis oils, manifested as gum formation. Gum and any insoluble substances that may be present in the pyrolysis oil can cause clogging problems in the process.

[0006] One method for removing these impurities from pyrolysis oil is to perform hydrotreating (HDT) in the presence of a catalyst. Such methods are described, for example, in WO2018 / 055555, WO2021 / 165178, or WO2022 / 144235.

[0007] Patent application WO2022 / 144235 describes a method for treating plastic pyrolysis oil, aimed at reducing and / or removing impurities contained in the pyrolysis oil to obtain an effluent compatible with a steam pyrolysis unit. The method includes the following steps: a) Hydrogenating the feedstock in the presence of at least hydrogen and at least one hydrogenation catalyst at an average temperature of 140 to 340°C, wherein the outlet temperature of step a) is at least 15°C higher than the inlet temperature of step a) to obtain a hydrogenated effluent. b) The hydrogenated effluent is hydrogenated in the presence of at least hydrogen and at least one hydrogenation treatment catalyst to obtain a hydrogenated effluent, wherein the average temperature of step b) is higher than the average temperature of step a). c) In the presence of an aqueous feed stream, the hydrotreated effluent is separated at a temperature of 50 to 370°C to obtain at least one gaseous effluent, an aqueous liquid effluent, and a hydrocarbon liquid effluent.

[0008] Therefore, one way to remove impurities from pyrolysis oil is to perform hydrotreating in the presence of a catalyst that is active in the sulfidated state.

[0009] In feedstocks containing plastic pyrolysis oil, the sulfur content of the feedstock is typically quite low. However, a minimum H2Spp is required in the hydrotreating reactor to maintain the catalyst in a sulfidated state, thus preventing its reduction. To maintain sufficient H2Spp in the reactor, and given the fact that the ex-plastic feedstocks do not contain sufficient sulfur, a sulfiding agent, typically DMDS (dimethyl disulfide), is usually, and in fact, continuously added to the feedstock. The sulfiding agent decomposes very rapidly at the reactor inlet through the action of temperature and hydrogen to produce H2S, thus providing the amount of H2S required to ensure a minimum and sufficient H2Spp.

[0010] The content of sulfur, oxygen, and / or nitrogen compounds typically depends on the source of the pyrolysis oil. Therefore, tire pyrolysis oils generally contain more impurities, especially sulfur compounds, than plastic and / or SRF pyrolysis oils.

[0011] The method according to the invention specifically utilizes the large amount of sulfur impurities contained in tire pyrolysis oil, using these impurities as a vulcanizing agent to maintain the sulfided state of the catalyst in the catalytic step of the method. Operation using tire pyrolysis oil thus significantly reduces the consumption of vulcanizing agents (typically DMDS type) that would normally be required. Invention Overview More specifically, the present invention relates to a method for processing a raw material comprising at least 5% by weight of tire pyrolysis oil relative to the total weight of the raw material, the tire pyrolysis oil having a sulfur content of 2,000 to 10,000 ppm by weight relative to the total weight of the tire pyrolysis oil, the method comprising the following steps: a) An optional selective hydrogenation step, carried out in a reaction section of a fixed-bed reactor employing at least one reactor having n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one selective hydrogenation catalyst, the selective hydrogenation reaction section being fed at least the feedstock and a hydrogen-containing gas stream, in the presence of at least one selective hydrogenation catalyst, at an average temperature of 100 to 280 °C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at spacetime velocities to obtain selective hydrogenation effluent; b) A hydrogenation step, which is carried out in a hydrogenation reaction section employing at least one fixed-bed reactor having n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one hydrogenation catalyst, the hydrogenation reaction section being fed at least the feedstock or the selective hydrogenation effluent obtained from step a) and a hydrogen-containing gas stream, the hydrogenation reaction section being at an average temperature of 140 to 400°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at spacetime to obtain hydrogenated effluent; c) A hydrotreating step, carried out in a hydrotreating reaction section employing at least one fixed-bed reactor having n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one hydrotreating catalyst, wherein when separation step d) is carried out between steps b) and c), the hydrotreating reaction section is fed at least the hydrocarbon effluent obtained from step d), and / or when separation step d) is carried out after step c), the hydrotreating reaction section is fed at least the hydrotreated effluent obtained from step b), and a hydrogen-containing gas stream, wherein the hydrotreating reaction section is maintained at an average temperature of 250 to 430°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at spacetime to obtain the hydrogenation treatment effluent; d) A separation step, performed between and / or after steps b) and c), wherein when separation step d) is performed between steps b) and c), the feed for this step includes a hydrogenated effluent obtained from step b), and / or when separation step d) is performed after step c), the feed includes a hydrogenated effluent obtained from step c), and an aqueous solution, wherein the step is performed in a separation section to obtain at least one gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

[0013] Therefore, the present invention relates to a method for purifying oil produced from tire pyrolysis, the oil optionally supplemented with pyrolysis oil obtained from plastics and / or SRF to remove at least some of its impurities, thereby enabling hydrogenation and thus enabling its upgrading by directly incorporating it into a fuel storage unit or by making it compatible with the processing of a steam cracking unit, while significantly reducing the consumption of vulcanizing agents.

[0014] Another advantage of this invention is the prevention of the risk of clogging and / or corrosion in the treatment unit performing the method of this invention, risks exacerbated by the presence of dienes, metals, and halogen compounds typically abundant in pyrolysis oils. These risks of clogging and / or corrosion are reduced in particular by a separation step involving washing with an aqueous solution. This washing / separation step is particularly effective in removing ammonium chloride salts formed by the reaction of chloride ions (released as HCl during the hydrogenation step, particularly by hydrogenation of chloride compounds, and subsequently dissolved in water) with ammonium ions (generated as NH3 during the hydrogenation treatment step by hydrogenation of nitrogen compounds, and / or supplied by injection of amines when chloride compound content is high, and subsequently dissolved in water). The high nitrogen compound content in tire pyrolysis oils reduces or even eliminates the need for nitrogen compound addition, which is another advantage of this invention.

[0015] Therefore, the method of the present invention can obtain hydrocarbon effluents from tire pyrolysis oil that at least partially remove impurities from the starting tire pyrolysis oil, thereby limiting operational problems that these impurities may cause, such as corrosion, coking or catalytic deactivation, particularly in steam cracking units and / or units located downstream of steam cracking units.

[0016] According to one variation, the raw material comprises plastics and / or solid recycled fuel pyrolysis oil.

[0017] According to one variant, the raw material consists of tire pyrolysis oil and plastic and / or (one or more) solid recycled fuel pyrolysis oil.

[0018] According to one variation, the raw material comprises tire pyrolysis oil at a content of 5% to 99% by weight relative to the total weight of the raw material and plastic and / or SRF pyrolysis oil at a content of 1% to 95% by weight relative to the total weight of the raw material.

[0019] According to one variant, separation step d) includes the following steps: d1) A separation step, wherein when separation step d) is carried out between steps b) and c), the feed includes a hydrogenated effluent obtained from step b), and / or when separation step d) is carried out after step c), the feed includes a hydrogenated effluent obtained from step c), said step being carried out at a temperature above the precipitation temperature of the ammonium halide and at a pressure substantially the same as that of step b), to obtain at least one first gaseous effluent and a liquid effluent, a portion of said liquid effluent optionally being recycled as a recirculated effluent upstream of step b); d2) A separation step, wherein the feed consists of a first gaseous effluent and at least a portion of the liquid effluent obtained from step d1), and an aqueous solution, said step being carried out at a temperature below the precipitation temperature of the ammonium halide and at a pressure substantially the same as or lower than that of step b), to obtain at least one second gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

[0020] According to one variation, step d1) is performed at a temperature of 200 to 450°C, and step d2) is performed at a temperature of 20°C or higher and below 200°C.

[0021] According to one variant, the method includes at least one step of pretreating the feedstock containing tire pyrolysis oil, the pretreatment step being carried out upstream of step a) and / or step b), and including an adsorption step and / or a filtration step and / or a centrifugation step and / or a sedimentation step and / or an electrostatic separation step and / or a washing step with an aqueous solution and / or a gas stripping step.

[0022] According to one variant, the method includes step e), fractionating all or part of the hydrocarbon effluent obtained from step d) when separation step d) is performed after step c), and / or all or part of the hydrotreated effluent obtained from step c) when separation step d) is performed between steps b) and c), to obtain at least one gas stream, a naphtha fraction, and at least one intermediate distillate fraction.

[0023] According to one variant, the method includes a hydrocracking step f), carried out in a hydrocracking reaction section employing at least one fixed-bed reactor having n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one hydrocracking catalyst. The hydrocracking reaction section is fed with at least a portion of the hydrocarbon effluent obtained from step d) when separation step d) is carried out after step c), and / or at least a portion of the hydrotreated effluent obtained from step c) when separation step d) is carried out between steps b) and c), and / or at least a portion of the middle distillate fraction obtained from fractionation step e), and a hydrogen-containing gas stream. The hydrocracking reaction section is maintained at an average temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹.-1 The process is carried out at a specific space-time velocity to obtain the first hydrocracking effluent.

[0024] According to one variant, the hydrogenation catalyst in step b) comprises a support selected from alumina, silica, aluminosilicate, magnesium oxide, clay, and mixtures thereof, and a hydrogenation-dehydrogenation functional containing at least one Group VIII element and at least one Group VIB element or at least one Group VIII element.

[0025] According to one variant, the hydrotreating catalyst in step c) comprises a support selected from alumina, silica, aluminosilicate, magnesium oxide, clay, and mixtures thereof, and a hydrotreating-dehydrogenating functional compound comprising at least one Group VIII element and / or at least one Group VIB element.

[0026] According to one variant, at least a portion of the hydrocarbon effluent obtained from step d) when separation step d) is carried out after step c), and / or at least a portion of the hydrotreated effluent obtained from step c) when separation step d) is carried out between steps b) and c), and / or at least one fraction obtained from fractionation step e) is fed to steam cracking step g), which is carried out in at least one cracking furnace at a temperature of 700 to 900°C and a relative pressure of 0.05 to 0.3 MPa.

[0027] According to one variant, the method includes a selective hydrogenation step (a).

[0028] According to one variant, the selective hydrogenation catalyst of step a) comprises a support selected from alumina, silica, aluminosilicate, magnesium oxide, clay and mixtures thereof, and a hydrogenation-dehydrogenation functional compound comprising at least one Group VIII element and / or at least one Group VIB element.

[0029] The present invention also relates to products obtained by the method according to the invention.

[0030] According to one variant, the product comprises, relative to the total weight of the product: • Total content of metallic elements less than or equal to 10.0 ppm by weight • Including iron content of 200 ppb or less by weight, and / or • Silicon content less than or equal to 5.0 ppm by weight, and / or • Sulfur content less than or equal to 100 ppm by weight, and / or • Nitrogen content less than or equal to 100 ppm by weight, and / or • Chlorine content less than or equal to 10 ppm by weight, and / or • Mercury content less than or equal to 5 ppb by weight.

[0031] According to one variant, the product contains 1% to 70% biocarbon according to ASTM D6866.

[0032] According to the present invention, the expressions "between... and..." and "to..." are equivalent and mean that the limit of the interval is included within the described numerical range. If this is not the case, and the limit is not included within the described range, the present invention provides such clarification.

[0033] In this specification, the term "comprising" is synonymous with "including" and "containing" and is inclusive or open-ended, and does not exclude other elements not mentioned. It should be understood that the term "comprising" includes the exclusive and closed term "consisting of".

[0034] In accordance with the present invention, various parameter ranges for a given step, such as pressure ranges and temperature ranges, can be used individually or in combination. For example, in accordance with the present invention, a preferred pressure range can be combined with a preferred temperature range.

[0035] According to the present invention, pressure is absolute pressure, also expressed as abs., and given in MPa absolute pressure (or MPa abs.), unless otherwise stated.

[0036] Specific and / or preferred embodiments of the invention may be described below. They may be implemented individually or in combination, and there is no limitation on the combination where technically feasible.

[0037] The following lists the chemical element groups according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, ed. DR. Lide, 81st edition, 2000-2001). For example, Group VIII (or VIIIB) according to the CAS classification corresponds to metals in columns 8, 9, and 10 according to the new IUPAC classification.

[0038] Metal content was measured using X-ray fluorescence.

[0039] Detailed description raw material According to the present invention, the raw material comprises at least 5% by weight of tire pyrolysis oil, said tire pyrolysis oil having a sulfur content of 2,000 to 15,000 ppm by weight relative to the total weight of the tire pyrolysis oil.

[0040] The feedstock may also contain any other pyrolysis oil, particularly plastic and / or solid recycled fuel (SRF) pyrolysis oil.

[0041] In addition, the feedstock may optionally be supplemented with conventional petroleum feedstocks and / or feedstocks derived from biomass conversion.

[0042] In the remainder of this document, the term "pyrolysis oil" is understood to mean oil produced from the pyrolysis of plastics and / or tires and / or SRF, unless otherwise stated. Where the source of the pyrolysis oil is significant, its source will be included (e.g., tire pyrolysis oil).

[0043] Tires are mainly composed of rubber (due to its elastic properties), (a mixture of cross-linked synthetic rubber and natural rubber type elastomers with additives such as silica, resin, sulfur, zinc oxide, and coke), as well as textiles and metal fibers (due to their reinforcing properties).

[0044] Plastic waste is typically a mixture of several polymers. This can be contained individually or as a mixture of (low and / or high density) polyethylene, polypropylene, polyethylene terephthalate, polyvinyl chloride, and polystyrene. In addition, depending on the intended use, plastics may contain other compounds besides polymers, such as plasticizers, pigments, dyes, or polymerization catalyst residues. Plastic waste may also contain small amounts of biomass, such as that derived from household waste.

[0045] Solid recycled fuel (SRF), also known as waste-derived fuel (RDF), is a solid, harmless waste prepared for energy upgrading, regardless of its origin from household and similar waste, economic activity waste, or construction and demolition waste. SRF is typically a mixture of any combustible waste, such as waste tires, food by-products (fat, animal meal, etc.), adhesive and wood waste, light fractions from shredders (e.g., from end-of-life vehicles, electrical and electronic equipment (WEEE), household and commercial waste), and residues from various types of waste recycling, including certain municipal waste, plastic waste, textiles, or wood. SRF often contains plastic waste.

[0046] The pyrolysis oil contained in the feedstock, whether derived from tires, plastics and / or SRF, comes from the step of pyrolyzing the feedstock containing tires, plastics and / or SRF in a pyrolysis unit.

[0047] The pyrolysis step can be carried out by thermal or catalytic pyrolysis, or by hydropyrolysis (pyrolysis in the presence of a catalyst and hydrogen).

[0048] The pyrolysis step is typically carried out at temperatures between 250°C and 750°C. The pyrolysis step can be performed under more or less harsh conditions.

[0049] The low-criticality pyrolysis step is carried out at a temperature of 250°C to 450°C, preferably 275°C to 425°C, and particularly preferably 300°C to 400°C. The low-criticality pyrolysis step produces pyrolysis oil rich in mono- and dienes, as well as a large amount of aromatics, and may include chlorinated compounds.

[0050] The high-severity pyrolysis step is carried out at a temperature of 450°C to 750°C, preferably 500°C to 700°C, and particularly preferably 550°C to 650°C. The high-severity pyrolysis step produces an aromatic-rich pyrolysis oil, which may include chlorinated compounds.

[0051] A pyrolysis unit may include one or more reactors configured to convert feedstock into gaseous and liquid products (e.g., simultaneously). The reactors (one or more) may contain one or more beds of inert material or pyrolysis catalysts, including sand, zeolite, or combinations thereof. Typically, the pyrolysis catalyst is capable of transferring heat to the components undergoing the pyrolysis process within the pyrolysis unit.

[0052] A pyrolysis unit may include one or more equipment items, such as one or more heated extruders, heated rotary kilns, heated tank reactors, vacuum heated vessels, enclosed heated surfaces (where feed flows along the walls), containers surrounded by one or more furnaces, or other equipment items that provide heated surfaces.

[0053] In one or more embodiments of the pyrolysis unit, purge gases are used in all or part of the pyrolysis stages to improve the pyrolysis of plastics and / or tires and / or SRF to produce valuable products, provide feedstock for vapor pyrolysis, or combinations thereof. The purge gases may contain hydrogen (H2), nitrogen (N2), vapor, product gases, or combinations thereof.

[0054] Pyrolysis oils, whether derived from tires, plastics, and / or SRF, typically have a boiling point range of 40°C to 1000°C, preferably 45°C to 650°C. Tire pyrolysis oils typically have a boiling point range of 100°C to 1000°C, preferably 150°C to 650°C. Plastic pyrolysis oils are generally lighter and typically have a boiling point range of 40°C to 700°C, preferably 45°C to 550°C.

[0055] It is also possible to use only partially pyrolyzed oil obtained through pyrolysis, such as pyrolyzed oil that has been pre-fractionated and / or tailed.

[0056] The density of pyrolysis oil, whether derived from tires, plastics, and / or SRF, typically ranges from 0.75 to 1.05 g / cm³, as measured at 15°C according to ASTM D4052. 3 The density is preferably 0.75 to 0.95 g / cm³. 3 The density of tire pyrolysis oil, measured at 15°C according to ASTM D4052 method, is typically 0.75 g / cm³. 3 Up to 1.05 g / cm 3 0.80 g / cm 3 Up to 0.98 g / cm 3The density of the plastic and / or SRF pyrolysis oil, measured at 15°C according to ASTM D4052 method, is typically low, and generally ranges from 0.75 to 0.99 g / cm³. 3 Preferably, the concentration is 0.75 to 0.95 g / cm³. 3 .

[0057] Pyrolysis oil, whether derived from tires, plastics, and / or SRF, is advantageously in liquid form at ambient temperature and particularly comprises a mixture of hydrocarbons, especially alkanes (n-alkanes and isoalkanes), alkenes (mono-alkenes and / or dienes), cycloalkanes, and aromatics. Specifically, depending on the source of the feedstock processed by the pyrolysis unit, the pyrolysis oil may contain up to 70% by weight of alkanes, up to 90% by weight of cycloalkanes, up to 90% by weight of alkenes, and up to 90% by weight of aromatics; it should be understood that the sum of alkanes, cycloalkanes, alkenes, and aromatics equals 100% by weight of hydrocarbons.

[0058] Tire pyrolysis oil typically contains a significant amount of aromatic compounds, for example, more than 30% by weight, preferably more than 40% by weight, and particularly preferably more than 50% by weight, relative to the weight of the tire pyrolysis oil.

[0059] Tire pyrolysis oils, whether derived from tires, plastics, and / or SRF, may contain dienes. Diolefin content is typically determined indirectly as a maleic anhydride value (MAV). This method is based on the Diels-Alder addition reaction of conjugated dienes with 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. MAV is expressed as milligrams (mg / g) of maleic anhydride reacting with 1 gram of sample. MAV in pyrolysis oils preferably varies between 5 and 100 mg / g.

[0060] Pyrolysis oils, whether derived from tires, plastics, and / or SRF, may contain partially bio-based compounds, such as tire pyrolysis oils produced from natural rubber-type elastomers, or plastic and / or SRF pyrolysis oils produced from plastic waste and / or SRF that may contain small amounts of biomass (e.g., derived from household waste). Biomass carbon content (according to C according to ASTM D6866) 14The isotope radiocarbon analysis method can be 0% to 70% by weight, preferably 0.1% to 60% by weight, relative to the total weight of the pyrolysis oil. In the case of tire pyrolysis oil, the latter may contain 20-70% by weight of biochar content (according to C according to ASTM D6866). 14 The method employs radiocarbon dating (i.e., 30% to 60% by weight, relative to the total weight of the tire pyrolysis oil). This allows for the incorporation of biological components into the product according to the method of the invention.

[0061] The impurity content in pyrolysis oils typically depends on the oil's source. Therefore, tire pyrolysis oils generally contain more impurities, particularly sulfur compounds, as well as nitrogen, oxygen, and halogen (especially chlorine) compounds, than plastic and / or SRF pyrolysis oils. They also typically contain more metals (especially silicon).

[0062] Pyrolysis oils, whether derived from tires, plastics, and / or SRF, may contain, and often do contain, impurities such as metals, especially iron and silicon, and halogen compounds, particularly chlorine compounds. These impurities may be present in high concentrations, for example, up to 500 or 700 ppm by weight contributed by halogen compounds, in fact even 1000 or 5000 ppm by weight of halogen elements (especially chlorine, as well as bromine, fluorine, or iodine), and typically in concentrations of 1 to 1000 ppm by weight, 1 to 700 ppm by weight, or 1 to 500 ppm by weight of halogen elements. Pyrolysis oils may contain up to 500 or 700 ppm by weight contributed by chlorine compounds, in fact even 1000 or 5000 ppm by weight of chlorine elements, and typically in concentrations of 1 to 1000 ppm by weight, 1 to 700 ppm by weight, or 1 to 500 ppm by weight of chlorine elements. Pyrolysis oil may contain up to 50 or 100 ppm by weight of bromine contributed by bromine compounds, and typically 1 to 100 ppm by weight or 1 to 50 ppm by weight of bromine.

[0063] Pyrolysis oils, whether derived from tires, plastics, and / or SRF, may contain up to 200 ppm by weight, and in fact even up to 1500 ppm by weight, of metallic or semi-metallic elements, typically ranging from 1 to 1500 ppm by weight or 1 to 200 ppm by weight. Alkali metals, alkaline earth metals, transition metals, post-transition metals, and metalloids can be classified in the same category as metallic contaminants and are referred to as metallic or metallic or semi-metallic elements. In particular, metallic or metallic or semi-metallic elements include silicon, iron, or both. Pyrolysis oils may contain up to 200 ppm by weight or 1000 ppm by weight of silicon, typically ranging from 1 to 1000 ppm by weight or 1 to 500 ppm by weight or 1 to 200 ppm by weight. Pyrolysis oils may also contain up to 50 ppm by weight or 100 ppm by weight of iron, typically ranging from 1 to 100 ppm by weight or 1 to 50 ppm by weight. Pyrolysis oils may also contain phosphorus, sodium, calcium, potassium, and magnesium.

[0064] Pyrolysis oil, whether derived from tires, plastics and / or SRF, may also contain other impurities, such as impurities contributed in particular by sulfur compounds, oxygen compounds and / or nitrogen compounds, typically in amounts below 40,000 ppm by weight, preferably below 15,500 ppm by weight, and typically in amounts between 1 and 40,000 ppm by weight or between 1 and 15,500 ppm by weight.

[0065] Oxygen compounds are typically present in concentrations of less than 15,000 ppm by weight, and preferably less than 10,000 ppm by weight, and typically between 1 and 15,000 ppm by weight or between 1 and 10,000 ppm by weight.

[0066] Nitrogen compounds are typically present in amounts below 10,000 ppm by weight, preferably below 8,000 ppm by weight, and generally in amounts of 1 to 10,000 ppm by weight or 1 to 8,000 ppm by weight. Tire pyrolysis oils typically contain more nitrogen compounds than plastic and / or SRF pyrolysis oils. Nitrogen compounds present in tire pyrolysis oils are typically present in amounts below 10,000 ppm by weight, preferably below 8,000 ppm by weight, and generally in amounts of 2,000 to 10,000 ppm by weight or 2,500 to 8,000 ppm by weight.

[0067] Nitrogen compounds present in plastic pyrolysis oil are typically present in amounts of less than 5,000 ppm by weight, and preferably less than 2,500 ppm by weight, and are typically between 5 and 5,000 ppm by weight or between 10 and 2,500 ppm by weight.

[0068] The sulfur content varies greatly depending on the source of the pyrolysis oil.

[0069] Sulfur compounds present in tire pyrolysis oil are typically present in amounts of less than 15,000 ppm by weight, and preferably less than 10,000 ppm by weight, and are typically between 2,000 and 15,000 ppm by weight or between 2,500 and 12,000 ppm by weight, preferably between 3,000 and 11,000 ppm by weight.

[0070] Sulfur compounds present in plastic pyrolysis oil are typically present in amounts of less than 2000 ppm by weight, and preferably less than 500 ppm by weight, and are typically between 5 and 2000 ppm by weight or between 10 and 500 ppm by weight.

[0071] Pyrolysis oil, whether derived from tires, plastics and / or SRF, may also contain other impurities such as heavy metals, such as mercury, arsenic, zinc and lead, for example up to 500 wt ppb of mercury or arsenic, and typically 1 to 300 wt ppb or 1 to 200 wt ppb of heavy metals.

[0072] The raw materials according to the method of the present invention comprise at least 5% by weight of tire pyrolysis oil relative to the total weight of the raw materials. The raw materials may consist solely of one or more types of tire pyrolysis oil.

[0073] The raw materials may also contain any other pyrolysis oils other than at least 5% by weight of tire pyrolysis oil, particularly plastic and / or solid recycled fuel (SRF) pyrolysis oils, in any proportion.

[0074] The raw material may consist solely of (one or more) tire pyrolysis oils and plastic and / or SRF pyrolysis oils. In this case, the raw material comprises 5% to 99% by weight, preferably 5% to 80% by weight, and particularly preferably 5% to 50% by weight, of tire pyrolysis oil, and 1% to 95% by weight, preferably 20% to 95% by weight, and particularly preferably 50% to 95% by weight, of plastic and / or SRF pyrolysis oil, relative to the total weight of the raw material.

[0075] The raw material may contain at least 50% by weight, preferably 70% to 100% by weight, of pyrolysis oil comprising at least 5% by weight of tire pyrolysis oil relative to the total weight of the raw material.

[0076] The raw materials for the method according to the invention may also include low amounts of conventional petroleum feedstocks or feedstocks derived from biomass conversion, other than tire pyrolysis oil and optionally plastic and / or SRF pyrolysis oil, typically 1% to 50% by weight of the feedstock, in fact even 1% to 30% by weight or 1% to 10% by weight, which are then co-processed with the pyrolysis oil of the feedstock.

[0077] Conventional petroleum feedstocks can advantageously be fractions or mixtures of naphtha, gas oil, or vacuum gas oil.

[0078] The feedstock obtained from biomass conversion can advantageously be selected from vegetable oils, algal oils or algal oils, fish oils, waste cooking oils, and fats of plant or animal origin, or mixtures of such feedstocks. The vegetable oils can advantageously be natural or wholly or partially refined, and can be obtained from plants selected from rapeseed, sunflower, soybean, palm, olive, coconut, coconut kernel, castor, cotton, peanut oil, flaxseed oil, and sea cabbage oil, as well as all oils obtained through genetic modification or hybridization from, for example, sunflower or rapeseed, this list is not limited to. The animal fats are advantageously selected from lard and fats derived from residues in the food industry or from the catering industry. Frying oils, various animal oils, such as fish oil, tallow, or lard, can also be used. The feedstock obtained from biomass conversion can also advantageously be selected from methyl esters of fatty acids of plant and / or animal origin or methyl esters of fatty acids of waste cooking vegetable oils.

[0079] Feedstocks derived from biomass conversion can also be selected from feedstocks derived from biomass through thermal or catalytic conversion methods, such as oils produced from biomass, particularly lignocellulosic biomass, via various liquefaction methods (such as hydrothermal liquefaction or pyrolysis). The term "biomass" refers to material derived from the most recently living organism, including plants, animals, and their byproducts. The term "lignocellulosic biomass" refers to biomass derived from plants or their byproducts. Lignocellulosic biomass consists of carbohydrate polymers (cellulose, hemicellulose) and aromatic polymers (lignin).

[0080] Raw materials obtained from biomass conversion can also be advantageously selected from raw materials obtained from the paper industry.

[0081] Preprocessing (optional) The feedstock containing tire pyrolysis oil may advantageously be pretreated in at least one optional pretreatment step prior to optional selective hydrotreating step a) and / or hydrotreating step b) to obtain a pretreated feedstock subsequently fed into step a) and / or step b).

[0082] According to one variant, this optional pretreatment step can reduce the amount of contaminants and solid particles. This optional pretreatment step is particularly effective in removing deposits that may form due to the unstable nature of pyrolysis oil and / or compatibility issues between two different feedstocks.

[0083] The optional pretreatment step can be carried out by any method known to those skilled in the art that can reduce the amount of contaminants. In particular, it may include adsorption and / or filtration and / or centrifugation and / or sedimentation and / or electrostatic separation and / or washing with aqueous solution and / or gas stripping.

[0084] The optional pretreatment steps are advantageously carried out at a temperature of 20 to 400°C, preferably 40 to 350°C, and a pressure of 0.15 to 10.0 MPa absolute pressure, preferably 0.2 to 7.0 MPa absolute pressure.

[0085] According to one variant, the optional pretreatment step is carried out in an adsorption section operating in the presence of at least one adsorbent. The adsorbent may be selected from zeolite, activated carbon, clay, silica, or alumina. Advantageously, the adsorbent contains less than 1% by weight of a metal element, and preferably contains no metal element. The metal element in the adsorbent should be understood to refer to elements of Groups VIB, VIIB, and VIII.

[0086] According to another variation, the optional pretreatment step is carried out in a washing section with an aqueous solution (e.g., water, an acidic or alkaline solution) or an organic solvent. This washing section may include equipment capable of contacting the feedstock with the aqueous solution and separating the phases to obtain, on the one hand, a pretreated feedstock, and on the other hand, an aqueous solution containing impurities. These equipment items may include, for example, stirred reactors, settling tanks, mixing-settling tanks, and / or co-current or counter-current scrubbing towers.

[0087] According to another variation, the optional pretreatment step is performed by filtration. The filtration step removes inorganic solids, sediments, and / or fine powders, particularly metals, metal oxides, and metal chlorides, from the feedstock. Filters with pore sizes (e.g., diameter or equivalent diameter) of less than 25 µm are typically used, preferably less than or equal to 10 µm, and even more preferably less than or equal to 5 µm. A series of filters with different pore sizes can also be used, particularly a series of filters with decreasing pore sizes along the feed circulation direction. These filter media are well known in industrial applications. For example, cartridge filters or self-cleaning filters are suitable.

[0088] According to one variant, filter aids, such as diatomaceous earth, can be used in filtration systems as seed crystals, i.e., suspended in the feed material, or as a pre-coating (coating) on ​​the filter.

[0089] Regardless of the type of filter, especially those with additives, two filters can be used in parallel so that one can be cleaned and maintained while the other is running.

[0090] According to another variation, the optional pretreatment step is carried out by centrifugation, sedimentation, or electrostatic separation.

[0091] According to another variation, the optional pretreatment step is carried out via gas stripping, thereby reducing the oxygen content in the feedstock. Gas stripping removes oxygen (O2) that may be dissolved in the feedstock, thus reducing the probability of free radical formation leading to polymerization in downstream steps. This method typically involves contacting the feedstock with an extraction gas (e.g., H2, N2, or a mixture thereof) to transfer at least a portion of the dissolved oxygen from the feedstock or fraction to the extraction gas, followed by separation of the extraction gas from the feedstock. Any dissolved H2 remaining in the feedstock after the gas stripping step is not a problem given the downstream hydrogenation step.

[0092] The optional preprocessing steps typically include one or more of the above-mentioned processes, preferably multiple processes.

[0093] The optional pretreatment step thus enables the preparation of pretreated feedstock for subsequent feeding into hydrogenation step b) and / or selective hydrogenation step a) (when the latter is present).

[0094] Selective hydrogenation step a) (optional) According to the invention, the method optionally includes a selective hydrogenation step a) carried out in a reaction section of a fixed-bed reactor having at least one fixed-bed reactor with n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one selective hydrogenation catalyst, the selective hydrogenation reaction section being fed at least the feedstock and a hydrogen-containing gas stream, in the presence of at least one selective hydrogenation catalyst, at an average temperature of 100 to 280°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.3 to 10.0 h⁻¹. -1 The process is carried out at a specific spacetime velocity to obtain a selective hydrogenation effluent.

[0095] The selective hydrogenation step is carried out under specific hydrogen pressure and temperature conditions, enabling the feedstock to be held in a liquid phase and possessing the amount of dissolved hydrogen required for selective hydrogenation of only the dienes present in the feedstock. The selective hydrogenation of dienes in the liquid phase thus avoids or at least limits the formation of "colloids," i.e., the polymerization of dienes and consequently the formation of oligomers and polymers, which could potentially clog the reaction sections of downstream steps b) and c). The selective hydrogenation step yields a selectively hydrogenated effluent, i.e., an effluent with reduced olefin content, particularly diene content, and preferably free of dienes.

[0096] The reaction section implements fixed-bed selective hydrogenation, advantageously at an average temperature of 100 to 280°C, preferably 110 to 250°C, and most preferably 130 to 200°C, in the presence of at least one selective hydrogenation catalyst, at an absolute pressure of 1.0 to 10.0 MPa, preferably 1.5 to 8.0 MPa, and very preferably 2.0 to 6.0 MPa, and a hydrogen partial pressure of 0.1 to 10.0 h⁻¹. -1Preferably, 0.2 to 5.0 h -1 And very preferably 0.3 to 3.0 h -1 The time-space velocity (HSV) is measured.

[0097] According to the present invention, the "average temperature" of the reaction section corresponds to the weighted average bed temperature (WABT), as is well known to those skilled in the art. The average temperature is advantageously determined based on the catalytic system used, the equipment, and its configuration. The average temperature (or WABT) is calculated as follows: WABT = (T 入口 + T 出口 ) / 2 Where T 入口 Temperature of the feed flow at the inlet of the reaction section, and T 出口 : Temperature of the effluent at the outlet of the reaction section. Unless otherwise specified, the "average temperature" of the reaction section is given under the conditions at the start of the cycle.

[0098] Hourly space velocity (HSV) is defined here as the ratio of the hourly volumetric flow rate of the feedstock to the volume of (one or more) catalysts.

[0099] The amount of hydrogen-containing gaseous feed (H2) fed into the reaction section of step a) advantageously results in a hydrogen coverage of 1 to 200 standard cubic meters of hydrogen (Sm) per cubic meter of feedstock. 3 / m 3 The preferred feedstock is 1 to 150 standard cubic meters of hydrogen (Sm). 3 / m 3 The preferred raw material is 5 to 100 standard cubic meters of hydrogen (Sm). 3 / m 3 ).

[0100] Hydrogen coverage is defined as the ratio of hydrogen volumetric flow rate to feed volumetric flow rate at 15°C, obtained under standard temperature and pressure conditions (expressed in standard cubic meters, denoted as Sm). 3 H2 / m 3 Raw materials), without considering the recycling portion.

[0101] The hydrogen-containing gas stream fed into the reaction section of step a) may consist of hydrogen supply and / or, in particular, recycled hydrogen obtained from the separation step d).

[0102] Advantageously, the reaction section of the selective hydrogenation step comprises 1 to 5 reactors. According to a particular embodiment of the invention, the reaction section comprises 2 to 5 reactors operating in a displaceable mode, referred to by the term PRS (Displaceable Reactor System) or "lead and lag". Combining at least two reactors in PRS mode allows for the isolation of one reactor to discharge spent catalyst, reloading the reactor with fresh catalyst, and bringing the reactor back into service without shutting down the process. PRS technology is specifically described in patent FR 2 681 871.

[0103] Advantageously, reactor internals, such as filter plate internals, can be used to prevent clogging of one or more reactors. Examples of filter plates are described in patent FR 3 051 375.

[0104] Advantageously, the at least one selective hydrogenation catalyst comprises a support, preferably a mineral support, and a hydrogenation-dehydrogenation functional compound.

[0105] The hydrotreating-dehydrogenating functional component particularly 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. The total content of oxides of Group VIB and VIII metal elements (i.e., the sum of Group VIB and VIII metal elements) is preferably from 1% to 40% by weight, and more preferably from 5% to 30% by weight, relative to the total weight of the catalyst. The weight ratio of Group VIB metal (or multiple metals) expressed as metal oxides to Group VIII metal (or multiple metals) is preferably from 1 to 20, and more preferably from 2 to 10. For example, the selective hydrotreating catalyst comprises, on a preferred mineral support, 0.5% to 10% by weight of nickel, preferably 1% to 5% by weight (expressed as nickel oxide NiO relative to the weight of the catalyst), and 1% to 30% by weight of molybdenum, preferably 3% to 20% by weight (expressed as molybdenum oxide MoO3 relative to the weight of the catalyst).

[0106] The support for the at least one selective hydrogenation catalyst is preferably selected from alumina, silica, aluminosilicate, magnesium oxide, clay, and mixtures thereof. The support may include doping compounds, particularly oxides selected from boron oxide, especially boron trioxide, zirconium oxide, cerium oxide, titanium oxide, phosphorus pentoxide, and mixtures of these oxides. Preferably, the selective hydrogenation catalyst comprises an alumina support, preferably doped with phosphorus and optionally with boron. When phosphorus pentoxide (P₂O₅) is present, its concentration is less than 10% by weight relative to the weight of alumina, advantageously at least 0.001% by weight relative to the total weight of alumina. When boron trioxide (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.

[0107] The selective hydrogenation catalyst is, for example, in extrusion form.

[0108] According to another aspect of the invention, the above-described selective hydrogenation catalyst further comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur. Such catalysts are commonly referred to by the term "additive catalyst". Typically, the organic compounds are selected from compounds containing one or more chemical functional groups selected from carboxylic acids, alcohols, thiols, thioethers, sulfones, sulfoxides, ethers, aldehydes, ketones, esters, carbonates, amines, nitriles, imides, oximes, ureas, and amides, or compounds including a furan ring, or sugars.

[0109] Preferably, the selective hydrogenation step may employ at least one guard bed upstream of one or more selective hydrogenation catalysts, containing an adsorbent of the type of alumina, silica-alumina oxide, zeolite, and / or activated carbon, said adsorbent optionally containing Group VIB and / or VIII metals from the selective hydrogenation stage. A series of guard beds with different particle diameters may also be used, particularly a series of guard beds with decreasing diameters along the feed circulation direction (also referred to as "gradation").

[0110] The selective hydrogenation effluent obtained at the end of the selective hydrogenation step has a lower content of impurities, particularly dienes, than the same impurities, particularly dienes, contained in the feedstock. Selective hydrogenation step a) converts at least 60%, preferably at least 70%, of the dienes contained in the feedstock. The selective hydrogenation effluent obtained at the end of the selective hydrogenation step is at least partially, and preferably entirely, preferably directly fed to hydrogenation step b).

[0111] Hydrogenation step b) According to the invention, the method includes a hydrogenation step b) carried out in a hydrogenation reaction section of a fixed-bed reactor having at least one hydrogenation reaction section having n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one hydrogenation catalyst, the hydrogenation reaction section being fed at least the feedstock or the selective hydrogenation effluent obtained from step a) and a hydrogen-containing gas stream, the hydrogenation reaction section being maintained at an average temperature of 140 to 400°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at a specific spacetime velocity to obtain a hydrogenated effluent.

[0112] Step b) is carried out, particularly under specific hydrogen pressure and temperature conditions, enabling the hydrogenation of dienes (possibly remaining after optional selective hydrogenation steps) and olefins at the beginning of the hydrogenation reaction section, while hydrodemetallization and hydrodechlorination are performed via a temperature rise curve, particularly at the end of the hydrogenation reaction section. The necessary amount of hydrogen is injected to enable the hydrogenation of at least some of the dienes and olefins present in the pyrolysis oil, to enable the hydrodemetallization of at least some metals (particularly silicon retention), and to convert at least some chlorine (to HCl). The hydrogenation of dienes and olefins thus avoids or at least limits the formation of "colloids," i.e., the polymerization of dienes and olefins, and therefore the formation of oligomers and polymers, which could clog the reaction section of the hydrotreatment step c). The hydrodemetallization, carried out in parallel with hydrogenation, particularly the silicon retention during step b), limits catalytic deactivation of the reaction section of the hydrotreatment step c). Furthermore, the conditions of step b) enable at least some chlorine conversion.

[0113] Those skilled in the art will readily understand that in hydrogenation step b), the above-described hydrogenation reaction is carried out, but some other hydrogenation treatment reactions are also carried out in parallel, and in particular hydrodesulfurization and hydrodenitrification reactions, even though these reactions are more advantageous in hydrogenation treatment step c), which is usually carried out at higher temperatures.

[0114] The temperature in step b), whether it is the average temperature (WABT), the reaction section inlet temperature, or the temperature rise between the reaction section inlet and outlet in step b), can be controlled, in particular, by injecting a diluent in step b), preferably by recycling a portion of the hydrocarbon effluent obtained from step d) (or d1) and / or at least a portion of one or more fractions obtained from step e), and particularly by the recycle ratio and / or the temperature of the recycle effluent.

[0115] The reaction section is carried out in the presence of at least one hydrogenation catalyst, advantageously at an average temperature (or WABT as defined below) of 140 to 400 °C, preferably 240 to 350 °C, and particularly preferably 260 to 330 °C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, preferably 1.5 to 8.0 MPa absolute, and 0.1 to 10.0 h⁻¹. -1 Preferably, 0.2 to 5.0 h -1 And very preferably 0.3 to 3.0 h -1 The time-space velocity (HSV) is measured.

[0116] The amount of hydrogen (H2) gas feed into the reaction section of step b) is advantageously such that the hydrogen coverage is 100 to 1500 standard cubic meters of hydrogen (Sm) per cubic meter of feed. 3 / m 3 The preferred feedstock is 200 to 1000 standard cubic meters of hydrogen (Sm). 3 / m 3 The preferred feedstock is 250 to 800 standard cubic meters of hydrogen (Sm). 3 / m 3 ).

[0117] The definitions of average temperature (WABT), HSV, and hydrogen coverage correspond to those described in step a) above.

[0118] Hydrogen can be obtained from fossil or renewable sources, such as from the gasification of plastic waste or produced by electrolysis.

[0119] Advantageously, the reaction section of step b) comprises 1 to 5 reactors, preferably 2 to 5 reactors, and particularly preferably two reactors. The advantage of a hydrogenation reaction section comprising multiple reactors is that it optimizes feedstock handling while reducing the risk of catalytic bed blockage (one or more), thereby avoiding unit shutdown due to blockage.

[0120] According to preferred variants, these reactors operate in a displaceable mode, referred to as a PRS (Displaceable Reactor System) or "lead and lag". Combining at least two reactors in PRS mode allows for the isolation of one reactor to discharge spent catalyst, reloading the reactor with fresh catalyst, and bringing the reactor back into service without shutting down the process. PRS technology is specifically described in patent FR 2 681 871.

[0121] According to a particularly preferred variant, the hydrogenation reaction section of step b) comprises two reactors operating in a displaceable mode.

[0122] The hydrogenation reaction section using at least one fixed-bed reactor can be operated in either a gas-liquid downflow or upflow manner.

[0123] Advantageously, reactor internals, such as filter plate type internals, can be used to prevent clogging of (one or more) reactors. Examples of filter plates are described in patent FR 3 051 375.

[0124] Advantageously, the hydrogenation catalyst comprises a support, preferably an inorganic support, and a hydrogenation-dehydrogenation functional compound.

[0125] According to one variant, the hydrotreating-dehydrogenating functional compound 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 from 1% to 40% by weight, more preferably from 5% to 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. The weight ratio of Group VIB metal (or multiple metals) to Group VIII metal (or multiple metals), expressed as metal oxides, is preferably from 1 to 20, and more preferably from 2 to 10.

[0126] According to this variant, the reaction section of step b) includes, for example, a hydrogenation catalyst comprising 0.5% to 12% by weight of nickel, preferably 0.9% to 10% by weight (in terms of nickel oxide NiO relative to the weight of the catalyst), on a preferred mineral support, preferably on an alumina support, and 1% to 30% by weight of molybdenum, preferably 3% to 20% by weight (in terms of molybdenum oxide MoO3 relative to the weight of the catalyst).

[0127] According to another variant, the hydro-dehydrogenation functional compound comprises, and preferably consists of, at least one Group VIII element, preferably nickel. According to this variant, the nickel oxide content is preferably from 1% to 50% by weight relative to the weight of the catalyst, more preferably from 10% to 30% by weight. This type of catalyst is preferably used in its reduced form, on a preferred mineral support, preferably on an alumina support.

[0128] The support for the hydrogenation catalyst is preferably selected from alumina, silica, aluminosilicate, magnesium oxide, clay, and mixtures thereof. The support may include doping compounds, particularly oxides selected from boron oxide, especially boron trioxide, zirconium oxide, cerium oxide, titanium oxide, phosphorus pentoxide, and mixtures of these oxides. Preferably, the hydrogenation catalyst comprises an alumina support, optionally 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 boron trioxide (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.

[0129] The hydrogenation catalyst is, for example, in extrusion form.

[0130] Very preferably, step b) may further employ at least one hydrogenation catalyst other than the hydrogenation catalyst(s) described above, wherein the at least one hydrogenation catalyst used in step b) comprises less than 1% by weight and at least 0.1% by weight, preferably 0.5% by weight, nickel (expressed as nickel oxide NiO relative to the weight of the catalyst) on an alumina support, and less than 5% by weight and at least 0.1% by weight, preferably 0.5% by weight, molybdenum (expressed as molybdenum oxide MoO3 relative to the weight of the catalyst). This catalyst with a low metal loading is preferably placed upstream or downstream of the aforementioned hydrogenation catalyst(s), preferably upstream.

[0131] Preferably, step b) may employ at least one guard bed upstream of one or more hydrogenation catalysts, comprising an adsorbent of the type of alumina, silica-alumina oxide, zeolite, and / or activated carbon, wherein the adsorbent optionally contains Group VIB and / or VIII metals. A series of guard beds with different particle diameters may also be used, particularly a series of guard beds with decreasing diameters along the feed circulation direction (also referred to as "gradation").

[0132] The hydrogenation step b) yields a hydrogenated effluent with reduced olefin content, particularly diene content, and reduced metal content, particularly silicon content, and halogen content, particularly chlorine content. Hydrogenation step b) typically converts at least 40%, preferably at least 60%, of the dienes and at least 40%, preferably at least 60%, of the olefins contained in the initial feedstock. The heat released through double bond saturation raises the temperature of the reaction medium and initiates the hydrogenation treatment reaction, particularly removing at least partially other contaminants, such as silicon and chlorine or nitrogen. Preferably, at least 50%, more preferably at least 75%, of chlorine and silicon are removed from the initial feedstock during step a). Typically, the silicon content in the hydrogenated effluent of step b) is less than 10 ppm by weight.

[0133] When separation step d) is performed after hydrogenation step c), the hydrogenated effluent obtained at the end of hydrogenation step b) is at least partially, and preferably entirely, sent directly to hydrogenation step c).

[0134] When separation step d) is carried out between steps b) and c), the hydrogenated effluent obtained at the end of hydrogenation step b) is at least partially and preferably all, preferably directly fed to separation step d).

[0135] Hydrogenation treatment step c) According to the invention, the processing method includes a hydrotreating step c) performed in a hydrotreating reaction section of a fixed-bed reactor having at least one hydrotreating reactor having n catalyst beds (n being an integer greater than or equal to 1), each catalyst bed containing at least one hydrotreating catalyst. When a separation step d) is performed between steps b) and c), the hydrotreating reaction section is fed at least the hydrocarbon effluent obtained from step d), and / or when a separation step d) is performed after step c), it is fed at least the hydrotreating effluent obtained from step b), and a hydrogen-containing gas stream. The hydrotreating reaction section is maintained at an average temperature of 250 to 430°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at a specific space-time velocity to obtain the hydrogenation treatment effluent.

[0136] Advantageously, step c) carries out hydrogenation reactions well known to those skilled in the art, and more particularly hydrogenation reactions such as aromatic hydrogenation, hydrodesulfurization, and hydrodenitrogenation. Furthermore, hydrogenation of the remaining halogenated compounds and olefins, as well as hydrodemetallization, continue.

[0137] The hydrotreating reaction section is advantageously carried out at a pressure comparable to that used in the hydrotreating step b), and typically at an average temperature higher than the average temperature used in the hydrotreating step b). Therefore, the hydrotreating reaction section is advantageously carried out at an average hydrotreating temperature of 250 to 430°C, preferably 280 to 380°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a duration of 0.1 to 10.0 h. -1 Preferably, 0.1 to 5.0 h -1 Preferably 0.2 to 2.0 h -1 Preferably 0.2 to 1 h -1 The procedure is carried out at a space-time velocity (HSV). The hydrogen coverage in step c) is advantageously 100 to 1500 standard cubic meters of hydrogen per cubic meter of fresh feed, and preferably 200 to 1000 standard cubic meters of hydrogen per cubic meter of fresh feed, more preferably 250 to 800 standard cubic meters of hydrogen per cubic meter of fresh feed. The definitions of mean temperature (WABT), HSV, and hydrogen coverage correspond to those described above.

[0138] Advantageously, the reaction section of step c) comprises 1 to 5 reactors, preferably 2 to 5 reactors. These reactors may be connected in series and / or in parallel and / or operated in a displaceable (or PRS) mode.

[0139] In a preferred embodiment, the hydrotreating reaction section comprises a single fixed-bed reactor containing n catalyst beds, where n is an integer greater than or equal to 1, preferably 1 to 10, and more preferably 2 to 5. In a particularly preferred embodiment, the hydrotreating reaction section of step b) comprises two reactors operating in a displaceable mode, followed by the hydrotreating reaction section of step c), which comprises a single fixed-bed reactor.

[0140] The hydrogenation reaction section using at least one fixed-bed reactor can be operated in either a gas-liquid downflow or upflow manner.

[0141] Advantageously, reactor internals, such as filter plate type internals, can be used to prevent clogging of (one or more) reactors.

[0142] Advantageously, the hydrotreating catalyst used in step c) may be selected from known hydrodemetallization, hydrotreating, or desiliconization catalysts, and combinations thereof, specifically for treating petroleum-based fractions. Known hydrodemetallization catalysts are, for example, those described in patents EP 0 113 297, EP 0 113 284, US 5 221 656, US 5 827 421, US 7 119 045, US 5622 616, and US 5 089 463. Known hydrotreating catalysts are, for example, those described in patents EP 0 113 297, EP 0 113 284, US 6 589 908, US 4 818 743, or US 6 332 976. Known silicon removal catalysts include, for example, those described in patent applications CN 102051202 and US 2007 / 080099.

[0143] In particular, the hydrotreating catalyst comprises a support, preferably a mineral support, and at least one metallic element with hydrotreating-dehydrogenating function. The metallic element with hydrotreating-dehydrogenating function 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. The total content of Group VIB and VIII metallic elements, expressed as oxides, is preferably from 0.1% to 40% by weight, more preferably from 5% to 35% 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. The weight ratio of Group VIB metals (or more metals) to Group VIII metals (or more metals), expressed as metal oxides, is preferably from 1.0 to 20, and more preferably from 2.0 to 10. For example, the hydrotreating reaction section includes a hydrotreating catalyst, which comprises 0.5% to 10% by weight of nickel, preferably 1% to 8% by weight (in terms of nickel oxide NiO relative to the total weight of the hydrotreating catalyst), on a mineral support, preferably on an alumina support, and 1.0% to 30% by weight of molybdenum, preferably 3.0% to 29% by weight (in terms of molybdenum oxide MoO3 relative to the total weight of the hydrotreating catalyst).

[0144] The support for the hydrotreating catalyst is advantageously selected from alumina, silica, aluminosilicate, magnesium oxide, clay, and mixtures thereof. The support may further comprise dopant compounds, particularly oxides selected from boron oxide, especially boron trioxide, zirconium oxide, cerium oxide, titanium oxide, phosphorus pentoxide, and mixtures thereof. 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 boron trioxide (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.

[0145] The hydrogenation catalyst is, for example, in extrusion form.

[0146] Advantageously, the hydrotreating catalyst of the method has a concentration of 200 to 800 m. 2 / g, preferably 250 to 600 m 2 / g, preferably 300 to 400 m 2 The specific surface area of ​​the hydrotreating catalyst was measured by the BET method, i.e., the specific surface area determined by nitrogen adsorption according to standard ASTM D 3663-78, which originates from the Brunauer-Emmett-Teller method described in The Journal of the American Chemical Society, 60, 309 (1938). Such a specific surface area can further improve the removal of contaminants, especially metals such as silicon.

[0147] According to another aspect of the invention, the above-described hydrogenation catalyst further comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur (such as the additive catalysts described above).

[0148] Preferably, step c) may employ at least one guard bed or a series of “staged” guard beds upstream of one or more hydrotreating catalysts, as described in step b) above.

[0149] Advantageously, the hydrotreating step c) enables at least 80%, preferably all, of the olefins and halogen compounds remaining after the hydrotreating step b), and at least partially converts other impurities present in the feedstock, such as aromatic compounds, metal compounds, sulfur compounds, nitrogen compounds, or oxygen compounds. Preferably, the nitrogen content at the outlet of step c) is less than 100 ppm by weight, and more preferably less than 10 ppm by weight. Preferably, the sulfur content at the outlet of step c) is less than 100 ppm by weight, and more preferably less than 10 ppm by weight. Step c) can also further reduce the content of contaminants, such as metal content, particularly silicon content. Preferably, the metal content at the outlet of step d) is less than 10 ppm by weight, and more preferably less than 2 ppm by weight, and the silicon content is less than 5 ppm by weight. Preferably, the halogen content at the outlet of step c) is less than 5 ppm by weight.

[0150] Catalyst preparation for selective hydrogenation, hydrogenation, hydrotreating, and hydrocracking steps (see below) is known and typically involves impregnating a Group VIII and Group VIB metal (where present) and optionally phosphorus and / or boron onto a support, followed by drying and then optionally calcination. In the case of additive catalysts, preparation is usually carried out simply by drying after the introduction of the organic compound, without calcination. The term "calcination" as used herein is understood to refer to heat treatment carried out in a gas containing air or oxygen at a temperature of 200°C or higher. Prior to use in the process steps, the catalyst is typically sulfided to form an active class.

[0151] Depending on the sulfur compound content in the initial feedstock to be treated, a stream containing a sulfurizing agent may be injected upstream of an optional pretreatment step, an optional selective hydrogenation step a) and / or a hydrogenation step b) and / or a hydrotreatment step c) and / or an optional hydrocracking step (when present), preferably upstream of a hydrogenation step b) and / or a hydrotreatment step c) to ensure sufficient sulfur content required for the formation of the catalytic active material (sulfided state).

[0152] The activation or sulfidation step is carried out by methods well known to those skilled in the art, and advantageously in a sulfur-reducing atmosphere in the presence of hydrogen and hydrogen sulfides. The sulfiding agent is preferably H₂S gas, elemental sulfur, CS₂, thiols, sulfides and / or polysulfides, hydrocarbon fractions with boiling points below 400°C containing sulfur compounds, or any other sulfur-containing compound used to activate the hydrocarbon feedstock to sulfidate the catalyst. The sulfur-containing compound is advantageously selected from alkyl disulfides, such as dimethyl disulfide (DMDS), alkyl sulfides, such as dimethyl sulfide, thiols, such as n-butyl mercaptan (or 1-butanethiol), and polysulfide compounds of the tert-nonyl polysulfide type. The catalyst can also be sulfided by the sulfur contained in the feedstock to be desulfurized. Preferably, the catalyst is sulfided in situ in the presence of the sulfiding agent and the hydrocarbon feedstock. Very preferably, the catalyst is sulfided in situ in the presence of a feedstock containing dimethyl disulfide.

[0153] The injection of vulcanizing agent is particularly necessary at the start of the catalytic cycle, when H2S is forming. Additional injections may be required throughout the catalytic cycle to compensate for natural losses. However, operating with tire pyrolysis oils with high sulfur loadings can significantly reduce vulcanizing agent consumption, or even eliminate its use entirely.

[0154] Separation step d) According to the present invention, the processing method includes a separation step d) performed between and / or after step b) and c), wherein when separation step d) is performed between steps b) and c), the feed for the step includes a hydrogenated effluent obtained from step b), and / or when separation step d) is performed after step c), the feed includes a hydrogenated effluent obtained from step c), and an aqueous solution, the step being performed in a separation section to obtain at least one gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

[0155] This washing / separation step is particularly effective in removing ammonium chloride salts formed by the reaction between chloride ions (released during the hydrogenation step, especially as HCl from the hydrogenation of chlorinated compounds, and subsequently dissolved in water) and ammonium ions (generated during the hydrogenation treatment step as NH3 from the hydrogenation of nitrogen compounds, and / or supplied by the injection of amines, and subsequently dissolved in water), thereby limiting the risk of blockage in the transmission lines and / or sections and / or transmission lines leading to the steam cracker in the method of the present invention due to ammonium chloride precipitation. This step also removes hydrochloric acid formed by the reaction of hydrogen ions dissolved in aqueous solution and halide ions released during hydrogenation and possible hydrogenation treatment steps, which are released from the hydrogenation of halogen compounds. When H2S is present in the hydrogenation treatment effluent obtained from step c), the washing / separation step can also remove ammonium sulfide ((NH4)2S) salts formed by the reaction between H2S (obtained from the hydrogenation desulfurization of sulfur compounds) and NH3 by dissolving it in aqueous solution.

[0156] The washing / separation step is advantageously carried out at a temperature of 50 to 450°C, preferably 100 to 440°C, more preferably 200 to 420°C, and particularly preferably 250 to 420°C. Operation within this temperature range is important (therefore, the hydrotreatment effluent should not be overcooled) to avoid blockage in the pipeline due to ammonium halide precipitation. Advantageously, the washing / separation step is carried out at a pressure close to that used in the hydrotreating step, preferably 1.0 to 10.0 MPa absolute, to facilitate hydrogen recirculation.

[0157] The aqueous solution can be water. It can also be an alkaline aqueous solution (e.g., by adding NaOH). Using an alkaline solution allows for the neutralization of hydrohalic acids and any dissolved salts.

[0158] The separation section may include any separation means known to those skilled in the art, particularly one or more flash tanks arranged in series, and / or one or more steam stripping and / or hydrogen stripping towers, and / or atmospheric distillation towers, and / or vacuum distillation towers. The washing / separation steps may advantageously be carried out in common or separate washing and separation equipment items, which are well known (separation tanks, pumps, heat exchangers, washing towers, etc., operable at various pressures and temperatures).

[0159] Advantageously, the washing / separation step includes injecting an aqueous solution, preferably water, into the upstream hydrogenated effluent and / or hydrogenated treated effluent of the washing / separation section to dissolve at least some and preferably all of the present hydrohalic acids (especially HCl) and any salts.

[0160] In an optional embodiment of the invention, the washing / separation step comprises injecting an aqueous solution into the hydrogenated effluent and / or the hydrotreated effluent, followed by a washing / separation section that advantageously includes a separation phase to obtain an aqueous effluent containing at least one hydrohalic acid (particularly HCl) and any dissolved salt, the hydrocarbon effluent, and a partially washed gaseous effluent. The aqueous effluent and the hydrocarbon effluent can then be separated in a settling tank to obtain the washed hydrocarbon effluent and the aqueous effluent. The partially washed gaseous effluent can be introduced in parallel into a washing tower, where it circulates countercurrently with the aqueous stream (preferably of the same nature as the aqueous solution injected with the hydrogenated and / or hydrotreated effluent), thereby removing at least some, and preferably all, of the hydrochloric acid contained in the partially washed gaseous effluent, and thus obtaining the gaseous effluent (preferably mainly containing hydrogen) and the acidic aqueous stream. The aqueous effluent obtained from the settling tank may optionally be mixed with the acidic aqueous stream, and optionally used as a mixture with the acidic aqueous stream in a water circulation loop to feed the washing / separation step upstream of the washing / separation section with the aqueous solution and / or with the aqueous stream in the washing tower. The water circulation loop may include the supply of water and / or alkaline solutions and / or an overflow capable of discharging impurities.

[0161] According to one embodiment, and depending on the chloride content in the feedstock, a nitrogen-containing compound (such as ammonia or amines, e.g., monoethanolamine, diethanolamine, monodiethanolamine, and / or aniline) stream can be injected upstream of the selective hydrogenation step and / or upstream of the hydrogenation step and / or the hydrotreating step. This ensures a sufficient amount of ammonium ions to combine with chloride ions formed during the hydrotreating step to form ammonium chloride salts, thereby limiting hydrochloric acid formation and thus limiting corrosion downstream of the separation section. The high nitrogen content in tire pyrolysis oil can reduce or even eliminate the need for nitrogen compound addition.

[0162] The aqueous effluent obtained at the end of the washing / separation step advantageously contains ammonium salts and / or hydrochloric acid.

[0163] The gaseous effluent obtained at the end of the washing / separation step advantageously contains hydrogen, preferably at least 60% by volume, and more preferably at least 70% by volume. The gaseous effluent obtained at the end of the washing / separation step contains very little chlorine, typically less than 3 ppm by weight, which allows it to be sent to a refining unit requiring hydrogen. According to one embodiment, the gaseous effluent can be at least partially recycled to one of the hydrogen-requiring steps (steps a), b), and c) of the method according to the invention, and optionally at least one hydrocracking step (if present). The recycling system may include a purification section (e.g., for adsorbing heavy metals such as mercury).

[0164] According to one variation, separation step d) is performed after hydrotreating step c). When separation step d) is performed after hydrotreating step c), the hydrocarbon effluent obtained from separation step d) can be partially or entirely fed directly to the steam cracking unit inlet or to an optional fractionation step e). Preferably, the hydrocarbon effluent is partially or entirely fed to fractionation step e). Preferably, separation step d) is performed after hydrotreating step c).

[0165] According to another variation, separation step d) is carried out between steps b) and c). When separation step d) is carried out between steps b) and c), the hydrocarbon effluent obtained from separation step d) is at least partially, and preferably entirely, preferably directly fed to the hydrotreating step c). This variation is advantageous when the feedstock contains a large amount of halogen compounds, particularly chlorine compounds. This is because the chlorine (in the form of HCl) contained in the effluent from hydrotreating step b) may deposit on the hydrotreating catalyst in step c), resulting in reduced activity and a shortened catalytic cycle time.

[0166] According to another variation, separation step d) can be performed after step c) and between steps b) and c). In this case, step d) is performed twice. This variation is advantageous for feedstocks containing large amounts of halogen compounds, especially chlorine compounds.

[0167] When separation step d) is performed after hydrotreating step c), at least a portion of the hydrocarbon effluent obtained from separation step d), and / or when separation step d) is performed between steps b) and c), at least a portion of the hydrotreating effluent obtained from step c) may be recycled upstream of step b) and / or upstream of step c) and / or upstream of selective hydrotreating step a). Recycling upstream of at least one reaction step advantageously allows for the dilution of impurities on the one hand, and on the other hand, for the control of the temperature in one or more reaction steps, where the reactions involved may be highly exothermic. The injection of the recycle stream may occur at the first catalyst bed level of the reaction section in steps a), b), and / or c), or between different catalyst beds in each section. When the reaction section comprises two reactors operating in a displaceable mode, at least a portion of the recycle stream may be recycled between the two reactors.

[0168] Advantageously, the amount of recycle effluent is adjusted such that the weight ratio between the recycle stream and the raw material (i.e., the raw material to be processed fed into the entire process) is less than or equal to 10, preferably less than or equal to 7, and 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 recycle stream is adjusted such that the weight ratio between the recycle stream and the raw material is from 0.01 to 10, preferably from 0.1 to 7, and particularly preferably from 0.2 to 5.

[0169] This recycling ratio is particularly effective in controlling the temperature rise in step b). This is because when the recycling ratio is high, the dilution rate of the feedstock is high, and the temperature rise at the beginning of the reaction section in step b), especially due to the diene hydrogenation reaction, can be controlled through the dilution effect.

[0170] Preferably, an additional hydrogen-containing gaseous feed stream is advantageously introduced at the inlet of each reactor (particularly in series operation) and / or at the inlet of each catalyst bed starting from the second catalyst bed in the reaction section. These additional gaseous feed streams are also called cooling feed streams. They are able to control the temperature in the reactor, where the reactions taking place are typically highly exothermic. The cooling feed stream can be injected into the reaction section of steps a), b), c), or f).

[0171] According to a preferred variant, at least a portion of the recirculated effluent may be advantageously cooled, or preheated if necessary, or kept at the same temperature at the outlet of separation step d) when separation step d) is performed after hydrotreating step c), and / or at the same temperature at the outlet of hydrotreating step c) when separation step d) is performed between steps b) and c).

[0172] According to one variation, the feedstock, preferably before being mixed with at least a portion of the recycled effluent, can be preheated to a temperature of up to 200°C, preferably up to 180°C, and particularly preferably up to 150°C by direct heating. Above this temperature, contact with the wall during direct heating may lead to the formation of gum and / or coke, which could cause scaling and increased pressure drop in the system heating the feedstock and in the catalyst bed(s) (especially when selective hydrogenation step a) is not performed). Heating the feedstock to a temperature above 150°C, preferably above 180°C, and particularly preferably above 200°C is preferably achieved through indirect heating with a thermal diluent (preferably at least a portion of the recycled effluent). Thus, raising the feedstock to a temperature above 150°C, preferably above 180°C, and particularly preferably above 200°C is achieved by mixing with a hotter liquid rather than by contact with a heated wall. This allows for the localization of high temperatures. This type of heating, which involves mixing with an inert hot liquid, can thus limit undesirable reactions such as diene polymerization (gel formation) and / or coke formation, and can adjust the inlet temperature of the feed stream in step b) so as to preferably initiate hydrogenation reactions of unsaturated bonds at the lowest possible temperature, while controlling the exothermic nature of these reactions through the dilution effect of the active substances.

[0173] In another variation, the feedstock is heated entirely through indirect heating of at least a portion of the recirculated effluent. In this case, the feedstock is not preheated before being mixed with at least a portion of the recirculated effluent.

[0174] Another heated feed stream advantageously consists of a hydrogen-rich gaseous effluent derived from the hydrogen supply and / or obtained from the gaseous effluent of separation step d).

[0175] Variation of the two-step separation step d) In one feasible embodiment of the invention, the separation step d), which may be performed after step c) and / or between steps b) and c), may include two steps: a first "hot" separation step followed by a second "cold" separation step.

[0176] This implementation scheme cleverly removes chlorine in the form of ammonium chloride salts through a combination of thermal separation step d1) followed by a cold separation / washing step d2). Chloride ions released as HCl by hydrogenation of chlorine compounds during steps b) and c) (hydrodechlorination), and ammonia generated as NH3 by hydrogenation of nitrogen compounds during step c) (hydrodenitrogenation), leave primarily as gaseous effluent through the thermal separation in step d1). This is because the high temperature of separation step d1) prevents the precipitation of ammonium halide salts formed by the reaction between chloride and ammonium ions. Separating the gaseous effluent and at least a portion of the liquid effluent at a lower temperature in step d2) causes these ammonium halide salts to precipitate. Washing with water in step d) dissolves these salts in the aqueous effluent. Thus, a chlorine-free hydrocarbon effluent is obtained.

[0177] According to this implementation scheme, separation step d) may include the following steps: d1) A separation step, wherein when separation step d) is performed between steps b) and c), the feed contains the hydrogenated effluent obtained from step b), and / or when separation step d) is performed after step c), the feed contains the hydrogenated effluent obtained from step c), said step being performed at a temperature above the ammonium halide precipitation temperature and at a pressure substantially the same as the pressure in step b), to obtain at least one first gaseous effluent and a liquid effluent, a portion of which may optionally be recycled as a recirculated effluent upstream of step b). d2) A separation step, wherein the feed consists of a first gaseous effluent and at least a portion of the liquid effluent obtained from step d1), and an aqueous solution, said step being carried out at a temperature below the ammonium halide precipitation temperature and at a pressure substantially the same as or lower than that of step b), to obtain at least one second gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

[0178] When separation step d) is performed after step c), and specifically between steps b) and c), a (first) "hot" separation step d1) (without step d2) can be performed, followed by (second) separation steps d1) and / or d2) after step c). Thus, at the first "hot" separation d1) between steps b) and c), at least one first gaseous effluent and liquid effluent are obtained, a portion of which is recycled as a recirculated effluent upstream of step b). Another portion of the gaseous and liquid effluents is then introduced into the hydrotreating step c). The hydrotreating effluent is then processed according to step d2) ("cold" separation with washing, optionally supplemented by step d1) after step c). The advantage of this configuration is that a hot recirculated stream is obtained, which can preheat the feedstock and control the exothermic reaction of step b), while salt removal requires only one washing operation (after step c).

[0179] Thermal separation step d1) The high temperature of separation step d1) avoids the precipitation of ammonium halide salts formed by the reaction of halide ions with ammonium ions, so as to recover the gaseous effluent containing most of the halides in the form of hydrohalic acid (HCl) and gaseous ammonia, and the liquid effluent containing very little halides and ammonia.

[0180] The term "precipitation temperature" for ammonium halide should be understood as the temperature at which gaseous ammonia and hydrohalic acid precipitate under given conditions (such as concentration and pressure), whether through reaction to form solid crystals of ammonium halide or through dissolution in water. According to thermodynamic principles, the precipitation temperature depends on the halide concentration and pressure. Under the conditions used in this method, the precipitation temperature of ammonium halide is typically between 150 and 300 °C, usually between 180 and 295 °C, and particularly between 200 and 290 °C.

[0181] The separation in step d1) must be carried out at a temperature higher than the precipitation temperature of the ammonium halide in order to recover the gaseous effluent containing most of the halide and gaseous ammonia in the form of hydrohalic acid (HCl), and the liquid effluent containing very little halide and ammonia. The separation in step d1) is typically carried out at a temperature of 200 to 450°C, preferably 220 to 330°C, and particularly preferably 240 to 300°C.

[0182] The term "pressure substantially equal to the pressure in step b)" should be understood to mean a pressure with a pressure difference of 0 to 1 MPa, preferably 0.005 to 0.3 MPa, and particularly preferably 0.01 to 0.3 MPa relative to the pressure in step b). Preferably, the pressure in step d1) is the pressure in step b) minus the pressure drop.

[0183] Separation step d1) can advantageously be carried out by any method known to those skilled in the art, such as a combination of one or more separators (tanks) and / or one or more stripping towers, wherein this or these separators (tanks) and / or towers may optionally be fed with stripping gas, such as a hydrogen-rich gas stream. Preferably, step d1) is carried out with a single separator (tank).

[0184] Separation step d1) is a separation step, referred to as a high-pressure or medium-pressure high-temperature separation step, and is also known by those skilled in the art as a HHPS (thermal high-pressure separator). Therefore, step d1) preferably employs a "thermal high-pressure" separator, with a pressure substantially equal to the operating pressure of step b).

[0185] Gas / liquid separation has an efficiency corresponding to solubility and Henry's law. This means that the equilibrium amount of halides in the form of hydrohalic acids will remain in the liquid effluent obtained in step d1). These halides will be released into the gaseous effluent during the cold separation step d2) and then discharged by dissolving in the aqueous effluent.

[0186] Advantageously, a portion of the liquid effluent from step d1) is recycled as a recirculated effluent to the upstream of step b) and / or optionally to the upstream of step a), the amount of the recirculated flow being as described above. In addition to the advantages of diluting impurities and temperature control described above, the recirculated flow is particularly capable of indirectly heating the feedstock or selective hydrogenation effluent by simply mixing the "cold" feedstock or "cold" selective hydrogenation effluent with the hot recirculated flow.

[0187] According to a preferred variant, at least a portion of the recirculated effluent may be advantageously cooled, or preheated if necessary, or maintained at the same temperature at the outlet of separation step d1) when separation step d) is performed after hydrotreating step c), and / or at the same temperature at the outlet of hydrotreating step c) when separation step d) is performed between steps b) and c).

[0188] The separation of high pressure and high temperature allows for maximizing energy recovery through the thermal recirculation of a portion of the liquid effluent. This is because the energy required to reach the inlet temperature of step b) and / or optionally step a) is at least partially provided by the heat obtained from a portion of the recirculated effluent in step d1), and also allows for reduction, and in fact even elimination, of optional preheating (which prevents gum formation by directly heating the feedstock to temperatures exceeding 200°C). Furthermore, the preferred operation of recirculating at least a portion of the liquid effluent under high pressure allows for energy savings in the pressurization required in step b).

[0189] Separation under high pressure and high temperature also allows for minimizing the amount of light fractions (naphtha fraction) contained in the liquid effluent. At this temperature, almost all of the light fractions (naphtha) of the effluent leave as a gaseous effluent, proceeding to the cold separation / washing step d2), while the liquid phase mainly consists of the heavy fractions (middle distillate fractions) of the feedstock. H2pp in step b) is advantageous when a portion of the liquid effluent is recycled to step b) because if the light fractions (naphtha) are not at least partially removed during the high-pressure and high-temperature separation, they may partially vaporize and reduce H2pp. The removal of light fractions containing naphtha can optionally be increased by slightly reducing the pressure upstream of at least one separator used in step d2), even though this use is not preferred due to the energy loss associated with the pressure reduction. Another option for increasing the removal of light fractions containing naphtha may include stripping, for example by injecting hydrogen-rich gas in step d1).

[0190] Cold separation and washing steps d2) According to the present invention, the processing method may include a separation step d2), wherein the feed consists of a first gaseous effluent and at least a portion of the liquid effluent optionally obtained from step d1), and an aqueous solution, the step being carried out at a temperature below the ammonium halide precipitation temperature and at a pressure substantially the same as or lower than the pressure of step b), to obtain at least one second gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

[0191] Step d2) involves separating the gaseous effluent and, optionally, at least a portion of the liquid effluent obtained from step d1) at a temperature below the precipitation temperature of ammonium halide, resulting in the precipitation of ammonium halide salts formed by the reaction of halide ions with ammonium ions particularly present in the gaseous effluent.

[0192] The cold separation step d2) feeds a gaseous effluent obtained solely from step d1), or a mixture of at least a portion of the gaseous effluent and liquid effluent obtained from step d1). Preferably, the cold separation step d2) feeds a mixture of at least a portion of the gaseous effluent and liquid effluent obtained from step d1). According to a variation, the cold separation step d2) feeds a mixture of gaseous effluent and all of the liquid effluent. Performing the cold separation step d2) on the mixture of gaseous and liquid effluents has the advantage of extracting not only contaminants from the gaseous effluent but also hydrohalic acid (HCl) dissolved in the liquid effluent. Specifically, the hydrohalic acid dissolved in the liquid effluent obtained from step d1) is released and also reacts with ammonia present in the gaseous effluent to form ammonium halides.

[0193] The washing with aqueous solution in step d2) allows these salts to dissolve in the aqueous effluent. This yields a hydrocarbon effluent free of halides, a gaseous effluent free of halides, and an aqueous effluent containing dissolved ammonium halide salts.

[0194] The separation temperature in step d2) should be below the precipitation temperature of ammonium halide to allow the ammonium halide salt to precipitate. The separation temperature in step d2) is above or equal to 20°C and below 200°C, preferably 25 to 120°C, and particularly preferably 30 to 70°C.

[0195] The term "pressure substantially equal to the pressure in step b)" should be understood to mean a pressure difference of 0 to 1 MPa, preferably 0.005 to 0.3 MPa, and particularly preferably 0.01 to 0.3 MPa, relative to the pressure in step b). Preferably, the pressure in step d2) is the pressure in step b) minus the pressure drop. Operating at least a portion of the separation step d2) at a pressure substantially equal to the operating pressure in step b) also facilitates the recirculation of hydrogen contained in the gaseous effluent.

[0196] Separation step d2) can also be performed at a lower pressure than step b).

[0197] Separation step d2) may also include a (first) separation step at a pressure substantially equal to the operating pressure of step b), followed by at least one other separation step at the same or lower temperature and at a pressure lower than that of each of the previous separation steps in step d2).

[0198] Separation step d2) can advantageously be carried out by any method known to those skilled in the art, such as a combination of one or more separators (tanks) and / or one or more stripping towers, which or these separators (tanks) and / or towers may optionally be fed with stripping gas, such as a hydrogen-rich gas stream. Separation step d2) may, for example, include a tower for stripping acidic water from the extracted aqueous effluent (also known as an acidic water stripping tower), a tower for washing the acidic gas to purify the hydrogen-rich gas before recirculation, and a tower for stabilizing the wash liquid effluent to remove dissolved gases.

[0199] Preferably, step d2) is carried out using a single separator (tank). Separation step d2) is preferably carried out in at least one separator tank, which is referred to as a high-pressure or medium-pressure cryogenic separator tank, also known by the name CHPS (Cold High Pressure Separator) to those skilled in the art. Therefore, step d2) preferably employs a "cold high-pressure" separator, with a pressure substantially equal to the operating pressure of step b).

[0200] Fractionation step e) (optional) The method according to the invention may include the following steps: when separation step d) is carried out after step c), fractionating all or part, preferably all, of the hydrocarbon effluent obtained from step d), and / or when separation step d) is carried out between steps a) and c), fractionating all or part, preferably all, of the hydrotreated effluent obtained from step c), to obtain at least one gaseous stream, a naphtha fraction, and at least one intermediate distillate fraction.

[0201] The term "naphtha fraction" should be understood to mean a hydrocarbon fraction containing compounds with boiling points generally below or equal to 175°C, particularly between 80°C and 175°C.

[0202] The term "middle distillate fraction" should be understood to mean a hydrocarbon fraction containing compounds with boiling points typically above 175°C. Heavy fractions may include middle distillate fractions, such as diesel and / or kerosene fractions. It may also contain heavier compounds.

[0203] Depending on the destination or use of the fraction obtained from fractionation step e), those skilled in the art will adjust the cut-off point during stripping and / or distillation operations. For example, it may be necessary to adjust the final boiling point of the naphtha fraction to 150, 175, or 200°C, or even 250°C.

[0204] Step e) specifically enables the removal of dissolved gases, such as ammonia, hydrogen sulfide, and light hydrocarbons having 1 to 4 carbon atoms, from the hydrotreated liquid effluent.

[0205] The optional fractionation step e) is advantageously carried out at an absolute pressure of less than or equal to 1.0 MPa, preferably from 0.1 to 1.0 MPa.

[0206] According to one embodiment, step e) can be carried out in a section of a stripping column that advantageously includes at least one reflux loop equipped with a reflux tank. Thus, the lightest compounds are entrained at the top of the column and enter the reflux loop containing the reflux tank, where gas / liquid separation occurs. The gas phase containing light hydrocarbons is drawn off from the reflux tank as a gaseous feed. The naphtha fraction is advantageously drawn off from the reflux tank as a liquid feed. The middle distillate fraction is advantageously drawn off from the bottom of the stripping column.

[0207] According to another embodiment, step e) can be carried out in a section comprising one or more separator tanks. Advantageously, at least one "low-pressure cryogenic" separator tank will be used to remove dissolved gases from the hydrotreating liquid effluent.

[0208] According to other implementation schemes, fractionation step e) may employ a stripping column or separator tank, followed by a distillation column, or only a distillation column.

[0209] Naphtha and middle distillate fractions, optionally mixed, may be sent wholly or partially to a steam cracking unit, where olefins may be (re)formed at the outlet to participate in polymer formation. Preferably, only a portion of the fractions is sent to the steam cracking unit; at least a portion of the remaining fraction is optionally recycled to at least one step of the process and / or sent to a fuel storage unit, such as a naphtha storage unit, diesel storage unit, or kerosene storage unit derived from conventional petroleum-based feedstocks.

[0210] According to the preferred embodiment, all or part of the naphtha fraction is sent to the steam cracking unit, while the middle distillate fraction is sent to the hydrocracking step f) and / or to the fuel storage unit.

[0211] In a particular embodiment, the optional fractionation step e) may result in, in addition to the gaseous feed stream, naphtha fractions (typically containing compounds with boiling points below or equal to 175°C, preferably 80 to 175°C), middle distillate fractions (typically containing compounds with boiling points above 175°C and below 385°C), and heavy hydrocarbon fractions (typically containing compounds with boiling points above or equal to 385°C). The naphtha fraction may be sent wholly or partially to a steam cracking unit and / or a naphtha storage unit derived from conventional petroleum feedstocks; it may also be recycled. The middle distillate fraction may also be sent wholly or partially to either a steam cracking unit, or to a diesel storage unit derived from conventional petroleum feedstocks, or, if present, to a hydrocracking step f), or also recycled. The heavy fraction, in itself, may be sent at least partially to a steam cracking unit, or, if present, to a hydrocracking step.

[0212] In another specific embodiment, the optional fractionation step e) can, in addition to the gaseous feedstream, yield a naphtha fraction (typically containing compounds with boiling points below or equal to 175°C, preferably 80 to 175°C), a kerosene fraction (typically containing compounds with boiling points above 175°C and below or equal to 280°C), a diesel fraction (typically containing compounds with boiling points above 280°C and below 385°C), and a heavy hydrocarbon fraction (typically containing compounds with boiling points above or equal to 385°C). The naphtha, kerosene, and / or diesel fractions can be sent, in whole or in part, either to a steam cracking unit, or to a naphtha, kerosene, or diesel pool derived from conventional petroleum feedstocks, or recycled. The diesel and / or kerosene fractions can also be sent, if present, to the hydrocracking step f). The heavy fractions, in themselves, can be sent, at least in part, to the steam cracking unit, or, if present, to the hydrocracking step f.

[0213] In another specific embodiment, the naphtha fraction obtained from step e) is fractionated into a heavy naphtha fraction (typically containing compounds with boiling points of 80 to 175°C) and a light naphtha fraction (typically containing compounds with boiling points below 80°C). At least a portion of the heavy naphtha fraction is fed to an aromatics complex including at least one naphtha reforming step to produce aromatic compounds. According to this embodiment, at least a portion of the light naphtha fraction is fed to the steam cracking step g) described below.

[0214] When the tire pyrolysis oil content in the feedstock is high (i.e., greater than 50% by weight of the feedstock), the naphtha fraction obtained from the method according to the present invention can also be sent to the catalytic reforming unit. This is because the high aromatic content in the tire pyrolysis oil enables the production of naphtha fractions with high aromatic and cycloalkanes content, which have high octane numbers (high RON / MON (RON = research octane number, and MON = motor octane number)).

[0215] The gaseous effluent (one or more) obtained from fractionation step e) may undergo additional purification and separation (one or more) to recover at least light hydrocarbons, particularly ethane, propane and butane, which may be advantageously fed, individually or as a mixture, to one or more furnaces in steam cracking step g) to increase the overall yield of olefins.

[0216] Hydrocracking step f) (optional) According to one variant, the method of the present invention may include step f), when separation step d) is performed after step c), hydrocracking at least a portion of the effluent obtained from step d), and / or when separation step d) is performed between steps b) and c), hydrocracking at least a portion of the hydrotreated effluent obtained from step c), and / or hydrocracking at least a portion of the middle distillate fraction obtained from fractionation step e).

[0217] Advantageously, step f) carries out a hydrocracking reaction well known to those skilled in the art, and more particularly enables the conversion of heavy compounds, such as those with boiling points above 175°C, into compounds with boiling points below or equal to 175°C. Other reactions, such as hydrogenation of alkenes and aromatics, hydrodemetallization, hydrodesulfurization, and hydrodenitrogenation, may continue.

[0218] Compounds with boiling points above 175°C contain more cycloalkanes, cycloalkanes-aromatics, and aromatics compared to lighter compounds, resulting in a higher C / H ratio. This high ratio is a cause of coking in steam crackers, thus requiring steam crackers dedicated to this fraction. When it is desirable to minimize the yield of these heavy compounds (middle distillate fractions) and maximize the yield of light compounds (naphtha fractions), these compounds can be at least partially converted to light compounds via hydrocracking, a fraction typically preferred by steam cracking units.

[0219] Therefore, the method of the present invention may include a hydrocracking step f), performed in a hydrocracking reaction section employing at least one fixed-bed reactor having n catalyst beds, where n is an integer greater than or equal to 1, each catalyst bed containing at least one hydrocracking catalyst, wherein when separation step d) is performed after step c), the feed to the hydrocracking reaction section contains at least a portion of the effluent obtained from step d), and / or when separation step d) is performed between steps b) and c), the feed contains at least a portion of the hydrotreated effluent obtained from step c), and / or the feed contains at least a portion of the middle distillate fraction obtained from fractionation step e), and a hydrogen-containing gas stream, wherein the hydrocracking reaction section is at an average temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at a specific space-time velocity to obtain the first hydrocracking effluent.

[0220] Therefore, the hydrocracking reaction section is advantageously operated at an average temperature of 250 to 480°C, preferably 320 to 450°C, at a hydrogen partial pressure of 1.5 to 20.0 MPa absolute, preferably 3 to 18.0 MPa absolute, and for 0.1 to 10.0 h. -1 Preferably 0.1 to 5.0 h -1 Preferably 0.2 to 4 hours -1 The procedure is carried out at a space velocity (HSV). The hydrogen coverage in step f) is advantageously 80 to 2000 standard cubic meters of hydrogen per cubic meter of feed to the fresh feedstock in step f), and preferably 200 to 1800 standard cubic meters of hydrogen per cubic meter of feed to the fresh feedstock in step f). The definitions of mean temperature (WABT), HSV, and hydrogen coverage correspond to those defined above.

[0221] Advantageously, the hydrocracking reaction section is carried out at a pressure comparable to that used in the reaction section of the hydrocracking step (b) or the hydrotreatment step (c).

[0222] Advantageously, step f) is carried out in a hydrocracking reaction section comprising at least one, preferably one to five, 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, each bed comprising at least one, and preferably no more than ten, hydrocracking catalysts.

[0223] The hydrocracking reaction section using at least one fixed-bed reactor can be operated in either a gas-liquid downflow or upflow manner.

[0224] The hydrocracking effluent can be recycled at least partially to steps b), c), d), or e).

[0225] The hydrocracking step can be performed in one step (step f) or in two steps (steps f) and f'). In the two-step process, the effluent from the first hydrocracking step f) is separated to obtain a hydrocarbon fraction (middle distillate) containing compounds with boiling points above 175°C. This fraction is introduced into a second hydrocracking step f'), which includes a dedicated second hydrocracking reaction section different from the first hydrocracking reaction section f). This configuration is particularly suitable when it is desired to produce only naphtha fractions.

[0226] The second hydrocracking step f') is carried out in a hydrocracking reaction section employing at least one fixed-bed reactor with n catalyst beds, where n is an integer greater than or equal to 1. Each catalyst bed contains at least one hydrocracking catalyst. The feed to the hydrocracking reaction section includes at least a portion of the first hydrocracking effluent obtained from the first hydrocracking step f') and a hydrogen-containing gas stream. The hydrocracking reaction section is maintained at an average temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at a specific space velocity to obtain the second hydrocracking effluent. The preferred operating conditions and catalysts used in the second hydrocracking step are those described in the first hydrocracking step. The operating conditions and catalysts used in the two hydrocracking steps may be the same or different.

[0227] The second hydrocracking step is preferably carried out in a hydrocracking reaction section, which includes at least one, preferably one to five, 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, each of the beds containing at least one, and preferably no more than ten, hydrocracking catalysts.

[0228] The operating conditions used in one or more hydrocracking steps typically enable a single-pass conversion of more than 15% by weight, and more preferably 20% to 95% by weight, to products with boiling points below or equal to 175°C, preferably below 160°C, and more preferably below 150°C. When the method is carried out in a two-step hydrocracking step, the single-pass conversion in the second step is kept moderate to maximize the selectivity of compounds (boiling points below or equal to 175°C, particularly 80 to below or equal to 175°C) in the naphtha fraction. The single-pass conversion is limited by using a high recycle ratio in the circulation loop of the second hydrocracking step. This ratio is defined as the ratio of the feed flow rate in step f) to the feed flow rate in step b); preferably, this ratio is 0.2 to 4, more preferably 0.5 to 2.5.

[0229] The hydrocracking effluent from the second hydrocracking step f') can be recycled at least partially to steps b), c), d), or e).

[0230] One or more hydrocracking steps therefore do not necessarily convert all hydrocarbons in the middle distillate fraction into hydrocarbons with boiling points below or equal to 175°C (naphtha fraction). After fractionation step e), there may therefore be a more or less significant proportion of compounds with boiling points above 175°C remaining. At least a portion of this unconverted fraction may be introduced into a second hydrocracking step f'). Another portion may overflow. Depending on the operating conditions of the method, the overflow may be 0% to 10% by weight, and preferably 0.5% to 5% by weight, relative to the fraction containing compounds with boiling points above 175°C entering the feedstock.

[0231] According to the present invention, one or more hydrocracking steps are carried out in the presence of at least one hydrocracking catalyst.

[0232] The hydrocracking catalysts used in one or more hydrocracking steps are conventional hydrocracking catalysts known to those skilled in the art, and are bifunctional types combining an acid functionalist with a hydro-dehydrogenation functionalist and optionally at least one binder matrix. The acid functionalist consists of a high surface area (typically 150 to 800 m²). 2 Supports exhibiting surface acidity ( / g) are provided, such as halogenated (especially chlorinated or fluorinated) alumina, combinations of alumina and boron oxide, amorphous silica-alumina oxides, and zeolites. Hydrogenation-dehydrogenation functionalizers are provided by at least one Group VIB metal and / or at least one Group VIII metal of the periodic table.

[0233] Preferably, the hydrocracking catalyst (one or more) comprises a hydrocracking-dehydrogenation functionalist, which includes at least one Group VIII metal selected from iron, cobalt, nickel, ruthenium, rhodium, palladium, and platinum, and preferably cobalt and nickel. Preferably, the catalyst (one or more) further comprises at least one Group VIB metal selected from chromium, molybdenum, and tungsten, alone or in mixtures, and preferably molybdenum and tungsten. Hydrocracking-dehydrogenation functionalists of the NiMo, NiMoW, or NiW type are preferred.

[0234] Preferably, the content of Group VIII metals in (one or more) hydrocracking catalysts is advantageously from 0.5% to 15% by weight, and more preferably from 1% to 10% by weight, the percentage being expressed as a 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.

[0235] Preferably, the content of a Group VIB metal in (one or more) hydrocracking catalysts is advantageously from 5% to 35% by weight, and more preferably from 10% to 30% by weight, the percentage being expressed as a 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.

[0236] (One or more) hydrocracking catalysts may optionally include at least one cocatalytic element deposited on the catalyst and selected from phosphorus, boron and silicon, optionally at least one Group VIIA element (preferably chlorine or fluorine), optionally at least one Group VIIB element (preferably manganese), and optionally at least one Group VB element (preferably niobium).

[0237] Preferably, the hydrocracking catalyst comprises at least one amorphous or low-crystallinity oxide-type porous mineral matrix selected from alumina, silica, aluminosilicate, alumina-boron oxide, magnesium oxide, silica-magnesium oxide, zirconium oxide, titanium oxide, or clay, alone or in mixtures, and preferably alumina or aluminosilicate, alone or in mixtures.

[0238] Preferably, the silicon-aluminum oxide contains more than 50% by weight of aluminum oxide, and more preferably more than 60% by weight of aluminum oxide.

[0239] Preferably, the hydrocracking catalyst (one or more) also optionally comprises zeolite selected from Y zeolite, preferably USY zeolite, alone or in combination with other zeolites (such as β zeolite, ZSM-12 zeolite, IZM-2 zeolite, ZSM-22 zeolite, ZSM-23 zeolite, SAPO-11 zeolite, ZSM-48 zeolite or ZBM-30 zeolite, alone or in mixtures). Preferably, the zeolite is USY zeolite alone.

[0240] When the catalyst includes zeolite, the zeolite content in (one or more) hydrocracking catalysts is advantageously from 0.1% to 80% by weight, preferably from 3% to 70% by weight, the percentage being expressed as a percentage of zeolite relative to the total weight of the catalyst.

[0241] Preferred catalysts include, and preferably consist of, at least one Group VIB metal and optionally at least one Group VIII non-precious metal, at least one co-catalyst element (and preferably phosphorus), at least one Y zeolite and at least one alumina binder.

[0242] Even more preferred catalysts include, and preferably consist of, nickel, molybdenum, phosphorus, USY zeolite and optional β zeolite and alumina.

[0243] Another preferred catalyst includes, and preferably consists of, nickel, tungsten, alumina and silicon aluminum oxide.

[0244] Another preferred catalyst includes, and preferably consists of, nickel, tungsten, USY zeolite, alumina, and silicon aluminum oxide.

[0245] The hydrocracking catalyst is, for example, in extrusion form.

[0246] In one variant, the hydrocracking catalyst used in the second hydrocracking step comprises a hydrocracking-dehydrogenation functional group containing at least one Group VIII noble metal selected from palladium and platinum, alone or in mixtures. The content of the Group VIII noble metal is advantageously from 0.01% to 5% by weight, and preferably from 0.05% to 3% by weight, expressed as a weight percentage of oxides (PtO or PdO) relative to the total weight of the catalyst.

[0247] According to another aspect of the invention, the hydrocracking catalyst described above further comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur as described above (additive catalysts).

[0248] Therefore, the hydrocarbon effluent or the hydrocarbon fraction(s) obtained by processing according to the method of the present invention has a composition compatible with the feedstock specifications at the inlet of the steam cracking unit. In particular, the composition of the hydrocarbon effluent or the hydrocarbon fraction(s) obtained from this method is preferably such that: • The total content of metallic elements is less than or equal to 10.0 ppm by weight, preferably less than or equal to 2.0 ppm by weight, more preferably less than or equal to 1.0 ppm by weight, and even more preferably less than or equal to 0.8 ppm by weight, wherein: The silicon (Si) content is less than or equal to 5.0 ppm by weight, preferably less than or equal to 1 ppm by weight, and even more preferably less than or equal to 0.6 ppm by weight. The iron (Fe) content is less than or equal to 200 ppb by weight. • Sulfur content less than or equal to 500 ppm by weight, preferably less than or equal to 200 ppm by weight. • The nitrogen content is less than or equal to 100 ppm by weight, preferably less than or equal to 50 ppm by weight, more preferably less than or equal to 10 ppm by weight, and particularly preferably less than or equal to 5 ppm by weight. • The total chlorine content is less than or equal to 5.0 ppm by weight, preferably less than 1.0 ppm by weight. • The content of olefin compounds (monoolefins and dienes) is less than or equal to 5.0% by weight, preferably less than or equal to 2.0% by weight, and more preferably less than or equal to 0.1% by weight. • The mercury content is less than or equal to 5 ppb by weight, preferably less than 3 ppb by weight.

[0249] Content is given as relative concentration by weight, percentage (%) by weight, parts per million (ppm) by weight, or parts per billion (ppb) by weight, relative to the total weight of the material flow under consideration.

[0250] Therefore, the method according to the invention enables the processing of pyrolysis oil to obtain at least one effluent that can be wholly or partially injected into at least one steam cracking unit.

[0251] Steam cracking step g) (optional) When separation step d) is performed after step c), at least a portion of the effluent obtained from step d), and / or when separation step d) is performed between steps b) and c), at least a portion of the hydrotreated effluent obtained from step c), and / or at least one of the liquid hydrocarbon fractions obtained from optional fractionation step e), may be partially or wholly sent to steam cracking step g).

[0252] Advantageously, the gaseous effluent containing ethane, propane, and butane obtained from separation step d) and / or fractionation step e) may also be sent, in whole or in part, to steam cracking step g).

[0253] The steam cracking step g) is advantageously carried out in at least one pyrolysis furnace at a temperature of 700 to 900°C, preferably 750 to 850°C, and at a relative pressure of 0.05 to 0.3 MPa. The residence time of the hydrocarbons is typically less than or equal to 1.0 second (s), preferably 0.1 to 0.5 s. Advantageously, steam is introduced upstream of the optional steam cracking step g) and after separation (or fractionation). The amount of water introduced is advantageously in steam form, preferably 0.3 to 3.0 kg of water per kg of hydrocarbons at the inlet of step g). Preferably, optional step g) is carried out in multiple pyrolysis furnaces in parallel so that the operating conditions can be adapted to the various feed streams to step g), particularly the stream obtained from step e), 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 group of furnaces operating in parallel. For example, the furnace may be dedicated to cracking middle distillate fractions.

[0254] The effluents from the various steam cracking furnaces are typically recombined prior to separation to form the effluent. It should be understood that the steam cracking step g) includes the steam cracking furnace and sub-steps related to steam cracking well known to those skilled in the art. These sub-steps may specifically include heat exchangers, columns, and catalytic reactors, and be recycled back to the furnace. Columns typically enable fractionation of the effluent to recover light fractions containing at least hydrogen and compounds having 2 to 5 carbon atoms, fractions containing pyrolysis gasoline, and optionally heavier fractions. Columns enable the separation of various components of the light fractions to recover at least ethylene-rich fractions (C2 fractions) and propylene-rich fractions (C3 fractions) and optionally butene-rich fractions (C4 fractions). Catalytic reactors particularly enable the hydrogenation of C2, C3, and even C4 fractions and pyrolysis gasoline. Saturated compounds, especially those having 2 to 4 carbon atoms, are advantageously recycled back to the steam cracking furnace to increase the overall yield of olefins.

[0255] The steam cracking step g) enables the production of an effluent containing at least one olefin having 2, 3, and / or 4 carbon atoms (i.e., C2, C3, and / or C4 olefins) in a satisfactory amount, particularly greater than or equal to 30% by weight, especially greater than or equal to 40% by weight, and in fact even greater than or equal to 50% by weight of total olefins having 2, 3, and / or 4 carbon atoms, relative to the weight of the steam cracking effluent in question. The C2, C3, and C4 olefins can then be advantageously used as polyolefin monomers.

[0256] Heavy metal adsorption steps (optional) Any gaseous effluent and / or any liquid effluent obtained from at least one of separation step d) or fractionation step e) may undergo an optional heavy metal adsorption step.

[0257] Optional adsorption steps enable the elimination or reduction of the amount of metallic impurities, particularly heavy metals such as arsenic, zinc, lead, and especially mercury, which may be present in the gaseous and liquid effluents. Metallic impurities, and especially heavy metals, are present in the feedstock. Some impurities, particularly mercury-based impurities, can be converted in one of the steps of the method according to the invention. Their converted forms are more easily captured. Their elimination or reduction may be particularly necessary when at least a portion of the gaseous and liquid effluents is intended to be fed directly or after undergoing one or more optional additional steps (such as fractionation step e) to a step with stringent specifications for metallic impurities (such as a steam cracking step).

[0258] Therefore, it is advantageous to perform an optional adsorption step on the gaseous effluents and / or hydrocarbon effluents obtained from the method according to the invention, particularly when at least one of these effluents or the raw materials contains more than 20 wt ppb, especially more than 15 wt ppb of heavy metal elements (As, Zn, Pb, Hg, etc.), and particularly when at least one of these effluents or the raw materials contains more than 10 wt ppb of mercury, more particularly more than 15 wt ppb of mercury.

[0259] The optional adsorption step is advantageously carried out at a temperature of 20 to 250°C, preferably 40 to 200°C, and at a pressure of 0.15 to 10.0 MPa absolute, preferably 0.2 to 1.0 MPa absolute.

[0260] The optional adsorption step can be carried out using any adsorbent known to those skilled in the art that can reduce the amount of such contaminants.

[0261] According to one variant, the optional adsorption step is carried out in an adsorption section operating in the presence of at least one adsorbent comprising a porous support and optionally at least one active phase, which may be based on elemental sulfur, or in the form of a metal sulfide or metal oxide, or also in the form of an elemental metal.

[0262] The porous support may be selected indiscriminately from alumina, aluminosilicate, silica, zeolite, and / or activated carbon. Advantageously, the porous support is based on alumina.

[0263] The adsorption section may include one or more adsorption towers. When the adsorption section includes two adsorption towers, one operating mode is "swing" operation, where one tower is online (operating) while the other is in standby. Another operating mode is with at least two towers operating in series in a switchable mode.

[0264] Preferably, the adsorption section includes an adsorption tower for gaseous effluents and an adsorption tower for liquid effluents.

[0265] Analytical methods used Analytical methods and / or standards for determining the characteristics of various material flows, particularly the characteristics of the raw materials and effluents to be treated, are known to those skilled in the art. These include, in particular, the information listed below. Other methods considered equivalent, particularly equivalent IP, EN, or ISO methods, may also be used: Table 1 (1) The MAV method is described in the paper: 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.

[0266] List of Attachments Figures 1 to 4 The details of the referenced elements enable a better understanding of the invention, rather than limiting it to... Figures 1 to 4 The specific implementation schemes described herein. The various implementation schemes presented can be used individually or in combination with each other, without any combination restrictions.

[0267] Figure 1 and Figure 2 An implementation scheme involving separation step d) performed after hydrogenation step c).

[0268] Figure 3 and Figure 4 An implementation scheme involving separation step d) between hydrogenation step b) and hydrogenation treatment step c).

[0269] Figure 1 A diagram illustrating a specific embodiment of the method of the present invention includes: • Selective hydrogenation step a) is carried out on feedstock 1 in the presence of hydrogen-rich gas 2, the feedstock comprising at least 5% by weight of tire pyrolysis oil relative to the total weight of the feedstock, the tire pyrolysis oil having a sulfur content of 2,000 to 10,000 ppm by weight relative to the total weight of the oil, to obtain selective hydrogenation effluent 3. • The selectively hydrogenated effluent 3 obtained from step a) is subjected to hydrogenation step b) in the presence of hydrogen-rich gas 4 and amines optionally supplied by feed stream 5 and sulfiding agents optionally supplied by feed stream 6 to obtain hydrogenated effluent 7. • Hydrotreating step c), the feed consists of the hydrogenated effluent 7 obtained from hydrotreating step b) and the hydrogen-rich gas stream 8, to obtain the hydrotreating effluent 9; • Separation step d), the feed consists of effluent 9 obtained from hydrotreatment step c), and is carried out in the presence of aqueous solution 12 to obtain at least one gaseous effluent 13, an aqueous effluent 14 and a hydrocarbon effluent 15.

[0270] Figure 2 A diagram illustrating another specific embodiment of the method of the present invention, which is based on Figure 1 The diagram includes step d), which is performed in two steps, namely steps d1) and d2), followed by a fractionation step e) and an additional hydrocracking step f).

[0271] (Optional) Selective hydrogenation steps a), hydrogenation steps b), and hydrogenation treatment steps c) as follows Figure 1 The separation step d), which is performed in two steps, specifically includes: • The hydrotreated effluent 9 is separated in step d1) under high pressure and high temperature (HHPS) to obtain at least one first gaseous effluent 10 and a liquid effluent 11, a portion 11a of which the liquid effluent 11 may be recycled to upstream of step b) or upstream of step a) (not shown). • Separation step d2) is carried out under high pressure and low temperature (CHPS) and the feed includes another portion 11b of the gaseous effluent 10 and liquid effluent obtained from step d1), as well as an aqueous solution 12, and makes it possible to obtain at least one hydrogen-containing gaseous effluent 13, an aqueous effluent 14 containing dissolved salts and dissolved H2S and NH3, and a hydrocarbon effluent 15. Optionally, the effluent 15 is subjected to a fractionation step e) to obtain at least one gaseous effluent 16, a naphtha fraction 17, and an intermediate distillate fraction 18.

[0272] At the end of step e), a portion of naphtha fraction 17 may be fed to a steam cracking process (not shown). Another portion of naphtha fraction 17 may be fed to selective hydrotreating step a), hydrotreating step b), and / or hydrotreating step c) (recycle stream, not shown).

[0273] exist Figure 2 In this process, at least a portion of the intermediate distillate fraction 18 obtained from step e) is fed to the hydrocracking step f), 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 may be recycled between separation steps d1) and d2), or further recycled upstream of separation step d) (not shown).

[0274] Figure 3 A diagram illustrating a specific embodiment of the method of the present invention includes: • Selective hydrogenation step a) is performed on feedstock 1, which contains at least 5% by weight of tire pyrolysis oil relative to the total weight of the feedstock, the tire pyrolysis oil having a sulfur content of 2,000 to 10,000 ppm by weight relative to the total weight of the oil, in the presence of hydrogen-rich gas 2, to obtain selective hydrogenation effluent 3. • The selectively hydrogenated effluent 3 obtained from step a) is subjected to hydrogenation step b) in the presence of hydrogen-rich gas 4 and amines optionally supplied by feed stream 5 and sulfiding agents optionally supplied by feed stream 6 to obtain hydrogenated effluent 7. • Separation step d), the feed consists of hydrogenated effluent 7 obtained from hydrogenation step b), and is carried out in the presence of aqueous solution 12 to obtain at least one gaseous effluent 13, an aqueous effluent 14 and a washed hydrocarbon effluent 23, which is particularly free of chlorine impurities. • Hydrotreating step c), the feed consists of a washed hydrotreating effluent 23, which is particularly free of chlorine impurities and obtained from separation step d), and a hydrogen-rich gas stream 8, to obtain hydrotreating hydrocarbon effluent 15.

[0275] Figure 4 A diagram illustrating another specific embodiment of the method of the present invention, which is based on Figure 3 The diagram includes step d), which is performed in two steps, namely steps d1) and d2), followed by a fractionation step e) and an additional hydrocracking step f).

[0276] (Optional) Selective hydrogenation step a) and hydrogenation step b) as follows Figure 3 The separation step d), which is performed in two steps, specifically includes: • The hydrogenated effluent 7 is separated in step d1) under high pressure and high temperature (HHPS) to obtain at least one first gaseous effluent 21 and a liquid effluent 22, a portion 22a of which the liquid effluent 22 may be recycled to upstream of step b) or upstream of step a) (not shown). • Separation step d2) is carried out under high pressure and low temperature (CHPS) and fed with another portion of the gaseous effluent 21 and liquid effluent 22b obtained from step d1), as well as an aqueous solution 12, and makes it possible to obtain at least one hydrogen-containing second gaseous effluent 13, an aqueous effluent 14 containing dissolved salts and dissolved H2S and NH3, and a washed hydrocarbon effluent 23. • Hydrotreating step c), the feed consists of washed hydrocarbon effluent 23, which is particularly free of chlorine impurities and obtained from separation step d), and hydrogen-rich gas stream 8, to obtain hydrotreated hydrocarbon effluent 15.

[0277] Optionally, the hydrotreated hydrocarbon effluent 15 is subjected to a fractionation step e) to obtain at least one gaseous effluent 16, a naphtha fraction 17, and an intermediate distillate fraction 18.

[0278] At the end of step e), a portion of naphtha fraction 17 may be fed to a steam cracking process (not shown). Another portion of naphtha fraction 17 may be fed to selective hydrotreating step a), hydrotreating step b), and / or hydrotreating step c) (recycle stream, not shown).

[0279] exist Figure 4 In this process, at least a portion of the intermediate distillate fraction 18 obtained from step e) is fed to the hydrocracking step f), 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 may be recycled between separation steps d1) and d2), or further recycled upstream of separation step d) (not shown).

[0280] Figures 1 to 4 Only the main steps and main feed streams are shown to facilitate a better understanding of the invention. It should be clearly understood that all equipment items required for operation are present (tanks, pumps, exchangers, ovens / furnaces, towers, etc.), even if not shown. It should also be understood that the hydrogen-rich gas feed stream (supply or recirculation stream) as described above can be injected at the inlet of each reactor or catalyst bed, or between two reactors or two catalyst beds. Apparatus for hydrogen purification and recirculation well known to those skilled in the art may also be employed. Example

[0281] Example 1 (according to the present invention) The feedstock processed in the method at a flow rate of 10,000 kg / h (10 T / h) is tire pyrolysis oil with the characteristics shown in Table 2 (i.e., containing 100% by weight of the tire pyrolysis oil).

[0282] Table 2: Raw Material Characteristics

[0283] The feedstock is mixed with the liquid recirculation stream (i.e., a portion of the liquid effluent from the hot tank obtained in separation step d) at a mass ratio of 80 / 20, i.e., 40 T / h of liquid recirculation stream for every 10 T / h of feedstock. The resulting mixture is then subjected to hydrogenation step b) in a fixed-bed reactor in the presence of hydrogen and a NiMo / alumina-type hydrogenation catalyst, under conditions shown in Table 3.

[0284] Table 3: Conditions for hydrogenation step b)

[0285] The conditions shown in Table 3 correspond to the conditions at the beginning of the loop.

[0286] At the end of hydrogenation step b), the observed conversion (= (initial concentration - final concentration) / initial concentration) is shown in Table 4.

[0287] Table 4: Degree of conversion of substances during hydrogenation step b)

[0288] The effluent from hydrotreating step b) is directly (without separation) fed to hydrotreating step b) in a fixed bed in the presence of hydrogen and a NiMo / alumina type hydrotreating catalyst, under the conditions shown in Table 5.

[0289] Table 5: Conditions for hydrotreating step c)

[0290] The conditions shown in Table 5 correspond to the conditions at the beginning of the loop.

[0291] The effluent from hydrotreating step c) undergoes separation step d), which includes separation steps d1) and d2). Separation step d1 involves a high-pressure separator tank operating at 300°C and 5.2 MPa absolute pressure. A fraction obtained from the bottom of the high-pressure tank allows a portion of the effluent to be recirculated to the inlet of hydrotreating step b) via a pump. A fraction obtained from the top of the high-pressure tank is mixed with a portion obtained from the bottom of the high-pressure tank that is not recirculated to the inlet of hydrotreating step b). This hydrocarbon mixture undergoes separation step d2), in which a water stream is injected. The hydrocarbon mixture and water stream are then treated in an acid gas scrubber and separator tank to obtain a gaseous fraction, an aqueous effluent containing salt, and a liquid effluent. The yields and masses of the various fractions obtained after separation are shown in Table 6 (yields correspond to the ratio of the mass of the various products obtained to the mass of the upstream feed in step b), expressed as a percentage and denoted as %m / m).

[0292] Table 6: Yields and masses of various products obtained after separation

[0293] The obtained liquid fraction can then be upgraded, either entirely or partially, in a steam cracking step to form olefins that can be polymerized to form recycled plastics.

[0294] The tire pyrolysis oil feedstock contains a significant amount of sulfur (4892 ppm by weight). This sulfur, existing in the form of sulfur molecules, is hydrogenated in the reaction section and converted into hydrogen sulfide (H₂S). This H₂S, present at a partial pressure of 0.019 MPa absolute at the hydrotreating reactor outlet (H₂Spp), helps maintain the sulfide phase of the NiMo / alumina catalyst, thereby preserving catalyst activity. Therefore, the catalytic activity loss in the hydrotreating step c) is low (<1 °C / month), and there is no need to inject hydrogen sulfide or sulfur compounds such as DMDS to maintain catalytic activity.

[0295] Example 2 (according to the present invention) The raw material processed in the method at a flow rate of 10,000 kg / h (10 T / h) is a 6 / 94 m / m mixture of tire pyrolysis oil and plastic pyrolysis oil from Example 1 (i.e., containing 6% by weight of the tire pyrolysis oil and 94% by weight of the plastic pyrolysis oil), which has the characteristics shown in Table 7.

[0296] Table 7: Raw Material Characteristics

[0297] The feedstock is mixed with the liquid recirculation stream (i.e., a portion of the liquid effluent obtained from the hot tank in separation step d) at a mass ratio of 80 / 20, i.e., 40 T / h of liquid recirculation stream for every 10 T / h of feedstock. The resulting mixture is then subjected to hydrogenation step b) in a fixed-bed reactor in the presence of hydrogen and a NiMo / alumina type hydrogenation catalyst, under conditions shown in Table 8.

[0298] Table 8: Conditions for hydrogenation step b)

[0299] The conditions shown in Table 8 correspond to the conditions at the beginning of the loop.

[0300] At the end of hydrogenation step b), the observed conversion (= (initial concentration - final concentration) / initial concentration) is shown in Table 9.

[0301] Table 9: Degree of conversion of substances during hydrogenation step b)

[0302] The effluent from the hydrogenation step b) is directly (without separation) fed to the hydrogenation step b) in a fixed bed in the presence of hydrogen and a NiMo / alumina type hydrogenation catalyst, under the conditions shown in Table 10.

[0303] Table 10: Conditions for hydrotreating step c)

[0304] The conditions shown in Table 10 correspond to the conditions at the beginning of the loop.

[0305] The effluent from hydrotreating step c) undergoes separation step d), which includes separation steps d1) and d2). Separation step d1 involves a high-pressure separator tank operating at 300°C and 5.2 MPa absolute pressure. A fraction obtained from the bottom of the high-pressure tank allows a portion of the effluent to be recirculated to the inlet of hydrotreating step b) via a pump. A fraction obtained from the top of the high-pressure tank is mixed with a portion obtained from the bottom of the high-pressure tank that is not recirculated to the inlet of hydrotreating step b). This hydrocarbon mixture undergoes separation step d2), in which a water stream is injected. The hydrocarbon mixture and water stream are then treated in an acid gas scrubbing tower and separator tank to obtain a gaseous fraction, an aqueous effluent containing salt, and a liquid effluent. The yields and masses of the various fractions obtained after separation are shown in Table 11 (yields correspond to the ratio of the mass of the various products obtained to the mass of the upstream feed in step b), expressed as a percentage and denoted as %m / m).

[0306] Table 11: Yields and masses of various products obtained after separation

[0307] The obtained liquid fraction can then be upgraded, either entirely or partially, in a steam cracking step to form olefins that can be polymerized to form recycled plastics.

[0308] The feedstock, a mixture of plastic pyrolysis oil and tire pyrolysis oil, has a significant sulfur content (355 ppm by weight). This sulfur, present in its sulfur-containing molecular form, is hydrogenated in the reaction section and converted to hydrogen sulfide (H₂S). This H₂S, present at a partial pressure of 0.0027 MPa absolute at the hydrotreating reactor outlet (H₂Spp), helps maintain the sulfided phase of the NiMo / alumina catalyst, thereby preserving catalyst activity. Therefore, the catalytic activity loss in the hydrotreating step c) is low (<1 °C / month), and no injection of hydrogen sulfide or sulfur compounds such as DMDS is required to maintain catalytic activity.

[0309] Example 3 (not in accordance with the present invention) According to Example 3, the raw material to be processed is a plastic pyrolysis oil having the characteristics shown in Table 12 (i.e., containing 100% by weight of the plastic pyrolysis oil).

[0310] Table 12: Raw Material Characteristics

[0311] The plastic pyrolysis oil according to Example 3 was treated under the same conditions as the tire pyrolysis oil according to Example 1.

[0312] The plastic pyrolysis oil feedstock contains very little sulfur (67 ppm by weight). This sulfur, existing in its sulfur-containing molecular form, is hydrogenated in the reaction section and converted to H₂S. This H₂S, in the form of a partial pressure of H₂S at an absolute pressure of 0.0004 MPa at the reactor outlet (H₂Spp), partially helps to maintain the sulfided phase of the NiMo / alumina catalyst. However, the H₂Spp obtained from this sulfur content (67 ppm by weight) in the feedstock is insufficient to maintain the sulfided phase of the catalyst throughout the cycle. This leads to rapid deactivation of the catalyst.

[0313] Example 4 (not in accordance with the present invention) According to Example 4, the feedstock to be treated is the same as the plastic pyrolysis oil to be treated in Example 3. The feedstock of Example 4 is treated under the same conditions as in Example 3, but H2S is added to the reaction system to achieve sufficient H2Spp. This H2S addition can be carried out by injecting dimethyl disulfide (DMDS) into the pyrolysis oil feedstock at the unit inlet. DMDS decomposes rapidly upon contact with the catalyst to produce CH4 and H2S, thereby generating sufficient H2Spp to maintain the sulfided form of the catalyst. This mode of operation results in high DMDS consumption, which is detrimental to the process economy.

[0314] Considering the sulfur content of the feedstock, the injected DMDS content was 800 ppm (equivalent to 544 ppm by weight of sulfur) to achieve the same hydrogen sulfide partial pressure as obtained in Example 2, i.e., 0.0027 MPa absolute pressure. For a feedstock capacity of 10 T / h, this represents a DMDS consumption of 8.0 kg / h, thus resulting in significant operational and environmental costs.

[0315] Example 5 (not in accordance with the present invention) According to Example 5, the raw material to be treated was the same as the plastic pyrolysis oil treated in Example 3, but it had a very high chloride content (2000 ppm by weight) and a nitrogen content of 700 ppm by weight. The plastic pyrolysis oil according to Example 5 was treated under the same conditions as the tire pyrolysis oil according to Example 3. Aniline was injected at a rate of 16.5 kg / h to remove all chloride ions.

[0316] Example 6 (according to the present invention) According to Example 6, the feedstock to be treated was a mixture comprising 15 wt% of the tire pyrolysis oil of Example 1 (nitrogen content of 2720 wt ppm) and 85 wt% of the plastic pyrolysis oil of Example 5 (sulfur content of 2000 wt ppm and nitrogen content of 700 wt ppm). The feedstock according to Example 6 was treated under the same conditions as the feedstock according to Example 5 (100 wt% plastic oil). No amine injection was required because the tire pyrolysis oil provided the nitrogen needed to remove all Cl ions. Furthermore, no injection of hydrogen sulfide or sulfur compounds such as DMDS was required to maintain catalytic activity.

Claims

1. A method for processing a raw material comprising at least 5% by weight of tire pyrolysis oil relative to the total weight of the raw material, the tire pyrolysis oil having a sulfur content of 2,000 to 10,000 ppm by weight relative to the total weight of the oil, the method comprising the following steps: a) An optional selective hydrogenation step, carried out in a reaction section of a fixed-bed reactor employing at least one reactor with n catalyst beds, where n is an integer greater than or equal to 1, each catalyst bed containing at least one selective hydrogenation catalyst, the selective hydrogenation reaction section being fed at least the feedstock and a hydrogen-containing gas stream, in the presence of at least one selective hydrogenation catalyst, at an average temperature of 100 to 280°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at a spacetime velocity to obtain a selective hydrogenation effluent. b) A hydrogenation step, which is carried out in a hydrogenation reaction section employing at least one fixed-bed reactor with n catalyst beds, where n is an integer greater than or equal to 1, each catalyst bed containing at least one hydrogenation catalyst, the hydrogenation reaction section being fed at least the feedstock or the selective hydrogenation effluent obtained from step a) and a hydrogen-containing gas stream, the hydrogenation reaction section being at an average temperature of 140 to 400°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at spacetime velocities to obtain hydrogenated effluent. c) A hydrotreating step, which is carried out in a hydrotreating reaction section employing at least one fixed-bed reactor with n catalyst beds, where n is an integer greater than or equal to 1, each catalyst bed containing at least one hydrotreating catalyst, wherein when separation step d) is carried out between steps b) and c), the hydrotreating reaction section is fed at least the hydrocarbon effluent obtained in step d), and / or when separation step d) is carried out after step c), the hydrotreating effluent obtained in step b) is fed at least the hydrogenated effluent, and a hydrogen-containing gas stream, wherein the hydrotreating reaction section is maintained at an average temperature of 250 to 430°C, a hydrogen partial pressure of 1.0 to 10.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at spacetime velocities to obtain the hydrogenation treatment effluent. d) A separation step, performed between and / or after step b) and step c), wherein when separation step d) is performed between step b) and step c), the feed for the step includes a hydrogenated effluent obtained from step b), and / or when separation step d) is performed after step c), the feed includes a hydrogenated effluent obtained from step c), and an aqueous solution, the step being performed in a separation section to obtain at least one gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

2. The method according to claim 1, wherein the raw material comprises plastic and / or solid recycled fuel pyrolysis oil.

3. The method according to any one of the preceding claims, wherein the raw materials consist of tire pyrolysis oil, plastic pyrolysis oil and / or solid recycled fuel pyrolysis oil.

4. The method according to any one of the preceding claims, wherein the raw materials comprise, by weight, 5% to 99% of tire pyrolysis oil and 1% to 95% of plastic pyrolysis oil and / or solid recycled fuel pyrolysis oil, in a proportion of 5% to 99% relative to the total weight of the raw materials.

5. The method according to any one of the preceding claims, wherein the separation step d) comprises the following steps: d1) A separation step, wherein when separation step d) is performed between steps b) and c), the feed contains the hydrogenated effluent obtained from step b), and / or when separation step d) is performed after step c), the feed contains the hydrogenated effluent obtained from step c), said step being performed at a temperature above the ammonium halide precipitation temperature and at a pressure substantially the same as the pressure in step b), to obtain at least one first gaseous effluent and a liquid effluent, a portion of said liquid effluent optionally being recycled as a recirculated effluent upstream of step b). d2) Separation step, wherein the feed consists of a first gaseous effluent and at least a portion of the liquid effluent optionally obtained from step d1), and an aqueous solution, said step being carried out at a temperature below the ammonium halide precipitation temperature and at a pressure substantially the same as or lower than that of step b), to obtain at least one second gaseous effluent, an aqueous effluent, and a hydrocarbon effluent.

6. The method according to claim 5, wherein step d1) is performed at a temperature of 200 to 450°C, and step d2) is performed at a temperature of 20°C or higher and lower than 200°C.

7. The method according to any one of the preceding claims, comprising at least one step of pretreating the raw material containing tire pyrolysis oil, said pretreating step being carried out upstream of step a) and / or step b), and including an adsorption step and / or a filtration step and / or a centrifugation step and / or a sedimentation step and / or an electrostatic separation step and / or a washing step with an aqueous solution and / or a gas stripping step.

8. The method according to any one of the preceding claims, comprising step e), when separation step d) is carried out after step c), fractionating all or part of the hydrocarbon effluent obtained from step d), and / or when separation step d) is carried out between steps b) and c), fractionating all or part of the hydrotreated effluent obtained from step c), to obtain at least one gas stream, a naphtha fraction and at least one intermediate distillate fraction.

9. The method according to any one of the preceding claims, comprising at least one hydrocracking step f), where n is an integer greater than or equal to 1, in a hydrocracking reaction section employing at least one fixed-bed reactor having n catalyst beds, each catalyst bed containing at least one hydrocracking catalyst, wherein when separation step d) is performed after step c), the hydrocracking reaction section is fed with at least a portion of the hydrocarbon effluent obtained from step d), and / or when separation step d) is performed between steps b) and c), the feed includes at least a portion of the hydrotreated effluent obtained from step c), and / or at least a portion of the middle distillate fraction obtained from fractionation step e), and a hydrogen-containing gas stream, wherein the hydrocracking reaction section is at an average temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa absolute, and a h⁻¹ of 0.1 to 10.0 h⁻¹. -1 The process is carried out at a specific space-time velocity to obtain the first hydrocracking effluent.

10. The method according to any one of the preceding claims, wherein the hydrogenation catalyst in step b) comprises a support selected from alumina, silica, aluminosilicate, magnesium oxide, clay and mixtures thereof, and a hydrogenation-dehydrogenation functional compound comprising at least one Group VIII element and at least one Group VIB element or at least one Group VIII element.

11. The method according to any one of the preceding claims, wherein the hydrogenation catalyst in step c) comprises a support selected from alumina, silica, aluminosilicate, magnesium oxide, clay and mixtures thereof, and a hydrogenation-dehydrogenation functional compound comprising at least one Group VIII element and / or at least one Group VIB element.

12. The method according to any one of the preceding claims, wherein when separation step d) is carried out after step c), at least a portion of the hydrocarbon effluent obtained from step d), and / or when separation step d) is carried out between steps b) and c), at least a portion of the hydrotreated effluent obtained from step c), and / or at least one of the fractions obtained from fractionation step e) is fed to steam cracking step g), which is carried out in at least one pyrolysis furnace at a temperature of 700 to 900°C and a relative pressure of 0.05 to 0.3 MPa.

13. The method according to any one of the preceding claims, comprising a selective hydrogenation step a).

14. The method according to the preceding claim, wherein the selective hydrogenation catalyst of step a) comprises a support selected from alumina, silica, aluminosilicate, magnesium oxide, clay and mixtures thereof, and a hydrogenation-dehydrogenation functional compound containing at least one Group VIII element and / or at least one Group VIB element.

15. The product obtained by the method according to any one of claims 1 to 14.

16. The product of claim 15, comprising, relative to the total weight of the product: • Total metal content less than or equal to 10.0 ppm by weight • Including iron content of 200 ppb or less by weight, and / or • Silicon content less than or equal to 5.0 ppm by weight, and / or • Sulfur content less than or equal to 100 ppm by weight, and / or • Nitrogen content less than or equal to 100 ppm by weight, and / or • Chlorine content less than or equal to 10 ppm by weight, and / or • Mercury content less than or equal to 5 ppb by weight.

17. The product according to claims 15 and 16, comprising 1% to 70% by weight of biocarbon content according to ASTM D6866.

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