PROCESS FOR TREATMENT OF PYROLYTIC OILS FROM PLASTICS AND / OR TIRES INCLUDING THE REMOVAL OF HALOGENIDES BY WASHING BEFORE A HYDROTREATMENT STAGE

Abstract

The present invention relates to a process for treating a pyrolysis oil from plastics and / or tires comprising: a) hydrogenating said feedstock in the presence of hydrogen and a catalyst to obtain a hydrogenated effluent, b) separating the hydrogenated effluent at a temperature above the precipitation temperature of the ammonium halides formed in step a) and at high pressure to obtain a gaseous effluent and a liquid effluent, c) separating at a temperature below the precipitation temperature and at high pressure and fed by the gaseous effluent and at least a part of the liquid effluent from step b) and an aqueous solution, to obtain at least a gaseous effluent, an aqueous effluent and a hydrocarbon effluent, d) hydrotreating said hydrocarbon effluent in the presence of hydrogen and a catalyst to obtain a hydrotreated effluent.

Description

Title of the invention: METHOD FOR TREATMENT OF PYROLYSIS OILS FROM PLASTICS AND / OR TIRES INCLUDING THE REMOVAL OF HALOGENIDES BY WASHING BEFORE A HYDROTREATMENT STEP technical field

[0001] The present invention relates to a process for treating pyrolysis oil from plastics and / or tires to obtain a hydrocarbon effluent that can be used in a gasoline, jet fuel, or diesel fuel storage unit or as feedstock for a steam cracking unit. More particularly, the present invention relates to a process for treating feedstock from the pyrolysis of plastic waste or tires to remove at least some impurities, and in particular halides, between two catalytic steps using hydrogen. Previous technique

[0002] Plastic waste is generally a mixture of several polymers, for example, mixtures of polyethylene, polypropylene, polyethylene terephthalate, polyvinyl chloride, and polystyrene. Furthermore, depending on their use, plastics may contain, in addition to polymers, other compounds such as plasticizers, pigments, dyes, or residues of polymerization catalysts. Plastic waste may also contain, in small quantities, biomass originating, for example, from household waste.

[0003] As for tires, they are mainly made of rubbers for their elastic properties (mixture of elastomers of the natural and synthetic cross-linked rubber type, with added additives of the silica type, resin, sulfur, zinc oxide, carbon black, etc.) and textile and metallic fibers for their reinforcing properties.

[0004] Plastics from collection and sorting channels or recycled tires can undergo a pyrolysis step to obtain, among other things, pyrolysis oils. These oils generally contain many impurities, in particular halogenated compounds, especially chlorine-based compounds, but also diolefins, olefins, metals, especially iron, silicon, or heteroelements such as sulfur, oxygen and nitrogen, and insolubles.

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

[0006] Another way of valorizing pyrolysis oils from plastics and / or tires is the use of these pyrolysis oils as a feed for a steam cracking unit in order to (re)create olefins, the latter being monomers that make up certain polymers.

[0007] The pyrolysis oils from plastics and / or tires can also be used as gasoline, jet or diesel fuels.

[0008] However, pyrolysis oils from plastics and / or tires often contain high levels of impurities that are incompatible with direct storage in a fuel storage unit or with steam cracking units or units located downstream of steam cracking units, particularly polymerization and selective hydrogenation processes. These impurities can cause operability problems, including corrosion (especially due to the presence of chlorine), coking, catalytic deactivation, and incompatibility issues with the target polymers. The presence of diolefins can also lead to pyrolysis oil instability, characterized by gum formation. Gums and any insolubles present in the pyrolysis oil can cause clogging problems in the processes.

[0009] One way to eliminate these impurities contained in the pyrolysis oils of plastics and / or tires is to carry out hydrotreatment (HDT) in the presence of catalysts.

[0010] The most problematic impurities found in plastic and / or tire oils are often halides, and more specifically chlorine. Indeed, chlorine is generally the limiting contaminant for treating pyrolysis oils in existing refinery units. Chlorine, even at low concentrations (e.g., <5 ppm by weight), is responsible for corrosion (in the form of HCl) that can occur in existing units whose metallurgy is generally not designed to withstand even low levels of chlorine. Another problem related to the presence of halides in pyrolysis oils, and particularly chlorine, is the formation of ammonium chloride salts. These salts are formed by a reaction between chloride ions, released by hydrodechlorination as HCl, and ammonium ions, generated by the hydrogenation of nitrogen compounds (hydrodeazotation) as NH3 during a hydrogenation and / or hydrotreating step.It is known that these ammonium chloride salts precipitate at a relatively low temperature (e.g., below 280°C), which creates clogging problems, particularly in transfer lines and / or in sections of a process downstream of hydrogenation / hydrotreating.

[0011] Processes for removing chlorine by hydrotreatment in a pyrolysis oil are known for example from documents: WO20020769, WO20016400, WO20239729, WO21105326 or WO16142809.

[0012] Document WO16142809 describes in particular a process including a step of pyrolysis, a hydrotreating step, a gas separation step (Cl to C4 containing H2S and HCl) and a C5+ liquid, a dechlorination step of the C5+ liquid by adsorption or a second hydrotreating step, possibly another gas separation step containing HCl, then a steam cracking step.

[0013] It is also known to remove the HCl formed during hydrotreating by washing with water after a hydrotreating step. Such washings are described in WO17083018, WO20254634, and WO22101333. None of these documents describes washing before a hydrotreating step, nor does it address the problem of deactivating the hydrotreating catalyst in a second hydrotreating step downstream of a first hydrotreating step by depositing chlorinated compounds.

[0014] Unpublished patent application FR 21 / 12.908 describes a process for treating a pyrolysis oil from plastics, comprising:

[0015] a) a step of hydrogenating the feed mixed with at least a portion of a liquid effluent from the separation step c) and hydrogen,

[0016] b) a hydrotreatment step of the effluent from step a) in the presence of hydrogen,

[0017] c) a separation step, fed by the effluent from step b), said step being operated at a temperature between 200 and 450°C and at a pressure substantially identical to the pressure of step b) to obtain at least a first gaseous effluent, and a liquid effluent, part of which is recycled upstream of step a),

[0018] d) a separation step, fed by the first gaseous effluent and another part of the liquid effluent from step c) and an aqueous solution, said step being operated at a temperature between 20 and below 200°C, and at a pressure substantially the same as or below the pressure of step c), to obtain at least a second gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

[0019] One of the objectives of application FR 21 / 12,908 is to remove chlorine in the form of ammonium chloride salts by combining hot separation step c) followed by cold separation / washing step d). Chloride ions, released by the hydrogenation of chlorinated compounds as HCl in steps a) and b) (hydrodechlorination), and ammonia generated by the hydrogenation of nitrogen compounds as NH3 in step b), in particular (hydrodeazotation), are largely removed in the gaseous effluent by the hot separation in step c). Indeed, the high temperature of this separation step c) prevents the precipitation of ammonium chloride salts, which are formed by the reaction between chloride and ammonium ions. The separation at a lower temperature in step d) of the gaseous effluent and part of the liquid effluent causes these ammonium chloride salts to precipitate.The water washing in this step d) allows these salts to dissolve in . the aqueous effluent. This yields a hydrocarbon effluent free of chlorine.

[0020] This same principle of chlorine removal by hot stripping to keep ammonia and halides in the gaseous phase until a quantity of water is available to collect the ammonium halides in solution is also described in WO22101333. In this document, this separation occurs, as in application FR 21 / 12.908, after two HDT steps in order to send the purified oil to a steam cracking step, the objective being to avoid the precipitation of solid ammonium halides on the internal surfaces of the equipment and corrosion.

[0021] The present invention provides an improvement of the process according to FR 21 / 12.908 by proposing the removal of chlorine not after the hydrotreatment step, but between the hydrogenation step and the hydrotreatment step.

[0022] The applicant observed that the chlorine contained in the feedstock is almost completely converted in the hydrogenation step. It also observed that by sending the effluent from the hydrogenation step directly into the hydrotreatment step, the chlorine (in the form of HCl) contained in the hydrogenation step effluent is deposited on the hydrotreatment catalyst, thus leading to a decrease in activity and a reduction in the duration of the catalytic cycle. In order to prevent clogging of the hydrotreatment catalyst and a decrease in activity, the present invention proposes removing the chlorine before the hydrotreatment step. Objectives and Summary of the Invention

[0023] More specifically, the invention relates to a process for treating a feed comprising a pyrolysis oil of plastics and / or tires comprising halogenated compounds, said process comprising:

[0024] a) a hydrogenation step carried out in a hydrogenation reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrogenation catalyst, said hydrogenation reaction section being fed at least by said feedstock optionally mixed with at least a portion of a liquid effluent from a separation step b) and a first gaseous stream comprising hydrogen, said hydrogenation reaction section being carried out at an average temperature between 140 and 400°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h*, to obtain a hydrogenated effluent,

[0025] b) a separation step, fed by the hydrogenated effluent from step a), said step being carried out at a temperature above the precipitation temperature of ammonium halides and at a pressure substantially identical to the pressure of step a) to obtain at least a first gaseous effluent and a liquid effluent of which part of it may be recycled upstream of step a),

[0026] c) a separation step, fed by the first gaseous effluent and optionally at least part of the liquid effluent from step b), and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of ammonium halides and at a pressure substantially the same as or lower than the pressure of step b), to obtain at least a second gaseous effluent, an aqueous effluent and a hydrocarbon effluent,

[0027] d) a hydrotreating step implemented in a hydrotreating reaction section, implementing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrotreating catalyst, said hydrotreating reaction section being fed at least by said hydrocarbon effluent from step c), optionally by at least a portion of the liquid effluent from step b), and a second gaseous stream comprising hydrogen, said hydrotreating reaction section being implemented at an average temperature between 250 and 430°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a hydrotreated effluent,

[0028] e) optionally a fractionation step of all or part of the hydrotreated effluent from step d), to obtain at least a third gaseous effluent, a naphtha cut and at least one middle distillates cut,

[0029] f) optionally a hydrocracking step carried out in a hydrocracking reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrocracking catalyst, said hydrocracking reaction section being fed by at least a portion of said hydrotreated effluent from step d) and / or by at least a portion of the middle distillate cut from step e) and a third gas stream comprising hydrogen, said hydrocracking reaction section being carried out at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly volumetric rate between 0.1 and 10.0 h 1 to obtain a first hydrocracking effluent.

[0030] The major advantage of the present invention is that it removes halogenated compounds before the hydrotreating step, thereby protecting the hydrotreating catalyst while preventing the precipitation of solid ammonium halides on the internal surfaces of the equipment and causing corrosion. Removing the halogenated compounds before the hydrotreating step thus prevents a loss of catalyst activity during the hydrotreating step and limits a reduction in the duration of the catalytic cycle.

[0031] The halogenated compounds contained in the charge are almost completely converted in step a) of hydrogenation into hydrogen halides (HCl, for example), as well as some of the nitrogen compounds into gaseous ammonia (NH3). Hot separation (step b), i.e., at a temperature above the precipitation temperature of ammonium halides, keeps the ammonia and halides primarily in the gaseous phase, resulting in a liquid effluent largely free of these contaminants. Cold separation (step c), i.e., at a temperature below the precipitation temperature of ammonium halides and in the presence of an aqueous solution, precipitates these ammonium halide salts while dissolving them in water. This yields a hydrocarbon effluent free of halogenated compounds, which is introduced into step d) of hydrotreating to remove the remaining impurities without the deposition of halogenated compounds on the catalyst.

[0032] Removing halogenated compounds before the hydrotreating step allows the hydrotreating step c) to be carried out at a lower average temperature than without prior removal of halogenated compounds, resulting in energy savings.

[0033] Furthermore, an increase in the cycle time of the hydrotreating section is observed. Cycle time refers to the duration of use of the catalyst without the need for catalyst replacement.

[0034] Step c) of cold separation can be carried out on the gaseous effluent (alone) or on the mixture of gaseous effluent and liquid effluent. Carrying out step c) of cold separation on the mixture of gaseous effluent and liquid effluent has the advantage of extracting not only contaminants from the gaseous effluent, but also hydrogen halides (HCl) dissolved in the liquid effluent.

[0035] Another advantage of the process according to the invention is to purify an oil from the pyrolysis of plastic waste of at least part of its impurities which makes it possible to hydrogenate it and thus to be able to valorize it in particular by incorporating it directly into the fuel storage unit or by making it compatible with a treatment in a steam cracking unit in order to be able to obtain in particular light olefins with increased yields which can be used as monomers in the manufacture of polymers.

[0036] Another advantage of the invention is to prevent the risks of clogging and / or corrosion of the treatment unit in which the process of the invention is implemented, the risks being exacerbated by the presence, often in large quantities, of diolefins, metals and halogenated compounds in the pyrolysis oil.

[0037] The process of the invention thus makes it possible to obtain a hydrotreated effluent from a pyrolysis oil that is at least partially free of impurities from the starting pyrolysis oil, thereby limiting operability problems, such as corrosion problems, coking or catalytic deactivation, which these impurities can generate, particularly in steam cracking units and / or in units located downstream of steam cracking units, including polymerization and hydrogenation units. Removing at least some of the impurities from the oils resulting from the pyrolysis of plastic waste also increases the range of applications for the target polymers, as incompatibilities in use are reduced.

[0038] According to one variant, step b) is carried out at a temperature between 200 and 450°C and step c) is carried out at a temperature greater than or equal to 20°C and less than 200°C.

[0039] According to one variant, step c) is fed by the first gaseous effluent and at least part of the liquid effluent from step b) and an aqueous solution.

[0040] According to one variant, at least part of the liquid effluent obtained in step b) is recycled upstream of step a).

[0041] According to one variant, the process includes the fractionation step e).

[0042] According to one variant, the process includes the hydrocracking step f).

[0043] According to one variant, in step a) the hydrogen cover is between 250 and 800 Nm3 of hydrogen per m3 of charge (Nm3 / m3).

[0044] According to one variant, the process includes a pretreatment step aO) of the pyrolysis oil fraction of plastics and / or tires, said pretreatment step being carried out upstream of step a) and includes an adsorption step and / or a filtration step and / or a centrifugation step and / or a decantation step and / or an electrostatic separation step and / or a washing step using an aqueous solution and / or a gas stripping step.

[0045] According to one variant, the hydrotreated effluent from the hydrotreatment step d), or at least one of the two liquid hydrocarbon cuts from step e), is sent in whole or in part to a steam cracking step g) carried out in at least one pyrolysis furnace at a temperature between 700 and 900°C and at a pressure between 0.05 and 0.3 MPa relative.

[0046] According to one variant, the reaction section of step a) implements at least two reactors operating in switchable mode.

[0047] According to one variant, a stream containing an amine and / or a sulfur compound is injected upstream of step a).

[0048] According to one variant, said hydrogenation catalyst comprises a support selected from alumina, silica, silica-aluminas, magnesia, clays and mixtures thereof and a hydro-dehydrogenating function comprising either at least one element from group VIII and at least one element from group VIB, or at least one element from group VIII.

[0049] According to one embodiment, said hydrotreating catalyst comprises a selected support in the group consisting of alumina, silica, silica-aluminas, magnesia, clays and their mixtures, and a hydro-dehydrogenating function comprising at least one element from group VIII and / or at least one element from group VIB.

[0050] According to one variant, the process further comprises a second hydrocracking step f') carried out in a hydrocracking reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrocracking catalyst, said hydrocracking reaction section being fed by at least one first hydrocracking effluent from the first hydrocracking step f) and a gas stream comprising hydrogen, said hydrocracking reaction section being carried out at a temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a second hydrocracking effluent.

[0051] According to one variant, said hydrocracking catalyst comprises a support selected from halogenated aluminas, combinations of boron and aluminum oxides, amorphous silica-aluminas and zeolites and a hydro-dehydrogenating function comprising at least one metal from group VIB selected from chromium, molybdenum and tungsten, alone or in mixture, and / or at least one metal from group VIII selected from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum.

[0052] The invention also relates to the product that can be obtained, and preferably obtained by the process according to the invention.

[0053] According to this variant, the product comprises, in relation to the total weight of the product: - a total content of metallic elements less than or equal to 10.0 ppm by weight, - of which an iron content less than or equal to 200 ppb by weight, - a silicon content less than or equal to 5.0 ppm by weight, - a sulfur content less than or equal to 500 ppm by weight, - a nitrogen content less than or equal to 100 ppm by weight, - a chlorine content less than or equal to 10 ppm by weight - a mercury content less than or equal to 5 ppb by weight.

[0054] According to the present invention, the pressures are absolute pressures, also noted as abs., and are given in absolute MPa (or abs. MPa), unless otherwise indicated.

[0055] In this description, the term "include" is synonymous with (means the same as) "include" and "contain," and is inclusive or open and does not exclude other elements not mentioned. It is understood that the term "include" includes the exclusive and closed term "consist."

[0056] In this description, the expression "between ... and ..." means that the limit values ​​of the interval are included in the range of values ​​described, unless otherwise specified.

[0057] In the sense of the present invention, the different parameter ranges for a given step, such as pressure ranges and temperature ranges, can be used alone or in combination. For example, in the sense of the present invention, a range of preferred pressure values ​​can be combined with a range of more preferred temperature values.

[0058] In the following, particular and / or preferred embodiments of the invention may be described. They may be implemented separately or in combination with each other, without limitation of combination where technically feasible.

[0059] In the following text, "pyrolysis oil" means an oil obtained from the pyrolysis of plastics and / or tires, unless otherwise indicated.

[0060] In the following, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUP AC classification.

[0061] The metal content is measured by X-ray fluorescence. Detailed description The charge

[0062] According to the invention, a "plastic pyrolysis oil" or "tire pyrolysis oil" is an oil, advantageously in liquid form at room temperature, obtained from the pyrolysis of plastics, preferably from plastic waste originating in particular from collection and sorting channels, or obtained from the pyrolysis of used tires.

[0063] It comprises in particular a mixture of hydrocarbon compounds, notably paraffins, mono- and / or di-olefins, naphthenes, and aromatics. At least 80% by weight of these hydrocarbon compounds preferably have a boiling point below 700°C, and preferably below 550°C. In particular, depending on the origin of the pyrolysis oil, it may comprise up to 70% by weight of paraffins, up to 90% by weight of naphthenes, up to 90% by weight of olefins, and up to 90% by weight of aromatics, it being understood that the sum of the paraffins, naphthenes, olefins, and aromatics is 100% by weight of the hydrocarbon compounds.

[0064] The oil may contain diolefins. The diolefin content is commonly determined indirectly as the maleic anhydride value (MAV). The method is based on the Diels-Alder addition reaction between conjugated diolefins and maleic anhydride. The method for determining the MAV is described in C. López-Garcia 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 mg of maleic anhydride reacted with 1 g of sample (mg / g). MAV varies between 5 and 100 mg / g in pyrolysis oils.

[0065] The density of the pyrolysis oil, measured at 15°C according to the ASTM D4052 method, is generally between 0.75 and 0.99 g / cm3, preferably between 0.75 and 0.95 g / cm3.

[0066] Pyrolysis oil may include, and most often does include, in addition to impurities such as metals, in particular iron, silicon, and halogenated compounds, in particular chlorinated compounds. These impurities may be present in the pyrolysis oil at high levels, for example up to 600 ppm by weight or even 700 ppm by weight or even 1000 ppm by weight and even 5000 ppm by weight of halogenated elements (in particular chlorine but also bromine, fluorine, iodine) supplied by halogenated compounds, and generally between 1 and 1000 ppm by weight or between 1 and 700 ppm by weight or between 1 and 600 ppm by weight of halogenated elements. Pyrolysis oil can contain up to 600 ppm by weight or even 700 ppm by weight or even 1000 ppm by weight and even 5000 ppm by weight of chlorine element supplied by chlorinated compounds, and generally between 1 and 1000 ppm by weight or between 1 and 700 ppm by weight or between 1 and 600 ppm by weight of chlorine element.

[0067] The oil may contain up to 200 ppm by weight, or even 1500 ppm by weight, of metallic or semi-metallic elements, and generally between 1 and 200 ppm by weight or between 1 and 1500 ppm by weight of metallic or semi-metallic elements. Alkali metals, alkaline earth metals, transition metals, post-transition metals, and metalloids can be considered as contaminants of a metallic nature, referred to as metallic or semi-metallic metals or elements. In particular, metallic or semi-metallic metals or elements include silicon, iron, or both. Pyrolysis oil may, in particular, contain up to 200 ppm by weight or even 1000 ppm by weight of silicon, and generally between 1 and 200 ppm by weight, or between 1 and 1000 ppm by weight, or between 1 and 500 ppm by weight of silicon. Pyrolysis oil may include up to 50 ppm by weight or 100 ppm by weight of iron, and generally between 1 and 50 ppm by weight or between 1 and 100 ppm by weight of iron.Pyrolysis oil may also contain phosphorus, sodium, calcium, potassium, and magnesium.

[0068] The pyrolysis oil may also include other impurities such as heteroatoms, notably sulfur compounds, oxygenated compounds, and / or nitrogen compounds, at levels generally below 40,000 ppm wt. of heteroatoms and preferably below 15,500 ppm wt. of heteroatoms, and generally between 1 and 40,000 ppm wt. or between 1 and 15,500 ppm wt. The sulfur compounds are generally present in a concentration less than 15000 ppm by weight and preferably less than 10000 ppm by weight, and generally between 1 and 15000 ppm by weight or between 1 and 10000 ppm by weight of sulfur compounds.

[0069] Oxygenated compounds are generally present in a content of less than 15000 ppm by weight and preferably less than 10000 ppm by weight, and generally between 1 and 15000 ppm by weight or between 1 and 10000 ppm by weight of oxygenated compounds.

[0070] Nitrogen compounds are generally present in a content of less than 10000 ppm by weight and preferably less than 5000 ppm by weight, and generally between 1 and 10000 ppm by weight or between 1 and 5000 ppm by weight of nitrogen compounds.

[0071] Pyrolysis oil may also include other impurities such as heavy metals like mercury, arsenic, zinc and lead, for example up to 100 ppb by weight or 200 ppb by weight of mercury or arsenic, and generally between 1 and 200 ppb by weight or between 1 and 100 ppb by weight of heavy metals.

[0072] The feed for the process according to the invention comprises at least one pyrolysis oil. Said feed may consist solely of pyrolysis oil(s). Preferably, said feed comprises at least 50% by weight, and more preferably between 70% and 100% by weight, of pyrolysis oil relative to the total weight of the feed, i.e., preferably between 50% and 100% by weight, and more preferably between 70% and 100% by weight of pyrolysis oil.

[0073] The feed of the process according to the invention may include, in addition to the oil or oils, a conventional petroleum feed or a feed from the conversion of biomass which is then co-treated with the pyrolysis oil of the feed.

[0074] The conventional petroleum feedstock can advantageously be a cut or a mixture of cuts of the naphtha, diesel or vacuum diesel type.

[0075] The feedstock resulting from biomass conversion may advantageously be selected from vegetable oils, algae or algal oils, fish oils, used cooking oils, and fats of vegetable or animal origin; or mixtures of such feedstocks. The oils / fats of vegetable and / or animal origin contain triglycerides and / or free fatty acids and / or esters. These vegetable oils may advantageously be crude or refined, wholly or partially, and derived from plants selected from rapeseed, sunflower, soybean, palm, olive, coconut, jatropha (purghera), copra, castor, cottonseed, peanut, linseed, and crambe oils, and all oils derived, for example, from sunflower or rapeseed by genetic modification or hybridization, this list not being exhaustive. Algal or fish oils are also suitable.Animal fats are advantageously chosen from lard and fats composed of residues from the food industry or from the catering industry. Frying oils, various animal oils such as fish oil, tallow, and lard can be used. can also be used.

[0076] The feedstock resulting from biomass conversion can also be selected from feedstocks obtained from thermal or catalytic biomass conversion processes, such as oils produced from biomass, particularly lignocellulosic biomass, using various liquefaction methods, such as hydrothermal liquefaction or pyrolysis. The term "biomass" refers to material derived from recently living organisms, including plants, animals, and their by-products. The term "lignocellulosic biomass" refers to biomass derived from plants or their by-products. Lignocellulosic biomass is composed of carbohydrate polymers (cellulose, hemicellulose) and an aromatic polymer (lignin).

[0077] The feed from biomass conversion can also advantageously be chosen from feeds from the paper industry.

[0078] Pyrolysis oil can be obtained from a thermal pyrolysis treatment, catalytic or can be prepared by hydropyrolysis (pyrolysis in the presence of a catalyst and hydrogen). (a0) Preprocessing step (optional)

[0079] Said feed comprising a pyrolysis oil can advantageously be pretreated in an optional pretreatment step aO), prior to step a) of hydrogenation, to obtain a pretreated feed which feeds step a).

[0080] This optional pretreatment step aO) makes it possible to reduce the quantity of contaminants and solid particles, in particular the quantity of iron and / or silicon and / or chlorine, which may be present in the feed containing pyrolysis oil. Thus, an optional pretreatment step aO) of the feed containing pyrolysis oil is advantageously carried out in particular when said feed contains more than 10 ppm by weight, in particular more than 20 ppm by weight, more particularly more than 50 ppm by weight of metallic elements and / or solid particles, and in particular when said feed contains more than 5 ppm by weight of silicon, more particularly more than 10 ppm by weight, or even more than 20 ppm by weight of silicon.Similarly, an optional pretreatment step aO) of the feed comprising a pyrolysis oil is advantageously carried out in particular when said feed comprises more than 10 ppm by weight, in particular more than 20 ppm by weight, more particularly more than 50 ppm by weight of chlorine.

[0081] Said optional pretreatment step aO) can be implemented by any method known to those skilled in the art that reduces the amount of contaminants. It may, in particular, include an adsorption step and / or a filtration step and / or a centrifugation step and / or a settling step and / or an electrostatic separation step and / or a washing step using a solution aqueous and / or a gaseous stripping step.

[0082] The optional pretreatment step aO) is advantageously carried out at a temperature between 20 and 400°C, preferably between 40 and 350°C, and at a pressure between 0.15 and 10.0 MPa abs, preferably between 0.2 and 7.0 MPa abs.

[0083] According to one embodiment, said optional pretreatment step aO) is carried out in an adsorption section operated in the presence of at least one adsorbent. The adsorbent can be chosen from a zeolite, activated carbon, clay, silica or alumina.

[0084] Advantageously, said adsorbent comprises less than 1% by weight of metallic elements, preferably being free of metallic elements. By metallic elements of the adsorbent, we mean the elements of groups 6 to 10 of the periodic table of elements (new IUP AC classification). The residence time of the charge in the adsorption section is generally between 1 and 180 minutes.

[0085] Said adsorption section of optional step aO) comprises at least one adsorption column, preferably comprising at least two adsorption columns, preferably between two and four adsorption columns, containing said adsorbent. When the adsorption section comprises two adsorption columns, one operating mode may be a so-called "swing" operation, in which one of the columns is online, i.e., in operation, while the other column is in reserve. When the adsorbent in the online column is depleted, that column is isolated while the reserve column is brought online, i.e., into operation. The depleted adsorbent can then be regenerated in situ and / or replaced with fresh adsorbent so that the column containing it can be brought online again once the other column has been isolated.

[0086] Another operating mode involves having at least two columns operating in series. When the absorbent in the leading column is depleted, this first column is isolated, and the spent absorbent is either regenerated in situ or replaced with fresh absorbent. The column is then placed back in the last position, and so on. This operation is called the permutable mode, or, in English, "PRS" for Permutable Reactor System, or "lead and lag" in the established English term. Combining at least two adsorption columns makes it possible to overcome the potential and potentially rapid poisoning and / or clogging of the adsorbent due to the combined action of metallic contaminants, diolefins, gums derived from diolefins, and insolubles that may be present in the pyrolysis oil being treated.The presence of at least two adsorption columns facilitates the replacement and / or regeneration of the adsorbent, advantageously without stopping the pretreatment unit, or even the process, thus reducing the risks of. clogging and therefore to avoid unit shutdown due to clogging, to control costs and to limit adsorbent consumption.

[0087] According to another embodiment, said optional pretreatment step aO) is carried out in a washing section with an aqueous solution, for example water or an acidic or basic solution. This washing section may include equipment for contacting the feed with the aqueous solution and separating the phases so as to obtain the pretreated feed on the one hand and the aqueous solution containing impurities on the other. This equipment may include, for example, a stirred reactor, a decanter, a mixer-decanter, and / or a co-current or counter-current washing column.

[0088] According to another embodiment, said optional pretreatment step aO) is implemented by filtration. The filtration step removes inorganic solids, sediments, and / or fines contained in the oil, including metals, metal oxides, and metal chlorides. A filter with a pore size (e.g., diameter or equivalent diameter) of less than 25 pm is generally used, preferably less than or equal to 10 pm, and even more preferably less than or equal to 5 pm. According to another embodiment, a filter with a pore size of less than 25 pm but greater than 5 pm may be used. Alternatively, a series of filters with different pore sizes may be used, in particular a series of filters with decreasing pore sizes in the direction of oil flow. These filter media are well known for industrial applications.Cartridge filters and self-cleaning filters are suitable examples. The dry extract can be measured, for example, by the Heptane Insolubles test, ASTM Method D-3279. The insolubles content in heptane must be reduced to less than 0.5% by weight, preferably less than 0.1%.

[0089] According to a particular embodiment, the pretreatment step aO) by filtration comprises at least one filter having a pore size of less than 10 microns, and preferably greater than 5 pm, optionally followed by a filtration system having a pore size of less than 2 pm and preferably less than 1 pm.

[0090] According to another particular embodiment, the pretreatment step aO) by filtration comprises at least one filter having a pore size of less than 10 pm, and preferably greater than 5 pm, followed by an electrostatic precipitation system.

[0091] According to another particular embodiment, the pretreatment step aO) by filtration comprises at least one filter having a pore size of less than 10 pm, and preferably greater than 5 pm, followed by a filter system(s) using filtration aids such as sand or diatomaceous earth.

[0092] According to another embodiment, said optional pretreatment step aO) is carried out by centrifugation. According to another embodiment, the pretreatment step aO) comprises centrifugation and filtration.

[0093] According to another embodiment, said optional pretreatment step aO) is carried out by decantation. According to another embodiment, the pretreatment step aO) comprises decantation and filtration.

[0094] According to another embodiment, said optional pretreatment step aO) is implemented by gas stripping, thereby reducing the oxygen content in the oil. Gas stripping can remove oxygen (O2) that may be dissolved in the feedstock, thereby reducing the probability of free radical formation leading to polymerization in downstream steps. The process generally involves contacting the oil with a stripping gas (e.g., H2, N2, or a mixture thereof), thereby transferring at least some of the dissolved oxygen from the oil to the stripping gas, followed by separation of the stripping gas from the oil. The volume of stripping gas relative to the volume of oil (both volumes measured under gas stripping conditions) is generally greater than 1, and preferably at least 3. In particular embodiments, the stripping gas may contain at least 60% (molar percent) H2.Any dissolved H2 remaining in the feed after the gas extraction step is not a problem, given the downstream hydrometallation / hydrotreating. Preferably, the gas extraction step is completed before any (pre)heating of the feed, in order to minimize potential fouling.

[0095] Said optional pretreatment step aO) generally comprises one or more, preferably several, treatments described above. It may, in particular, comprise a sequence of a washing step using an aqueous solution and / or an adsorption step, followed by a gas stripping step, followed by a filtration step and / or a centrifugation step. All these steps are preferably carried out before any (pre)heating of the feed.

[0096] Said optional pretreatment step aO) may also optionally be fed by at least a portion of the liquid effluent from step b) of the process and / or a portion of at least one of the fractions from step e), either mixed with or separately from the feed containing pyrolysis oil. Recycling at least a portion of the liquid effluent from step b) notably increases sedimentation and thus, after possible filtration, improves the pretreatment of the feed.

[0097] Said optional pretreatment step aO) thus makes it possible to obtain a pretreated feed which then feeds the hydrogenation step a). (a) Hydrogenation step

[0098] According to the invention, the process comprises a step a) of hydrogenation carried out in a hydrogenation reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrogenation catalyst, said hydrogenation reaction section being fed at least by said feedstock, possibly pretreated, possibly mixed with at least a portion of liquid effluent from step b) and a first gaseous stream comprising hydrogen, said hydrogenation reaction section being operated at an average temperature between 140 and 400°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a hydrogenated effluent.

[0099] Step a) is carried out in particular under hydrogen pressure and temperature conditions that allow the hydrogenation of diolefins and olefins at the beginning of the hydrogenation reaction section, while a rising temperature profile allows hydrodemetallation and hydrodechlorination, particularly at the end of the hydrogenation reaction section. A sufficient quantity of hydrogen is injected to allow the hydrogenation of at least some of the diolefins and olefins present in the plastic pyrolysis oil, the hydrodemetallation of at least some of the metals, including the retention of silicon, and also the conversion of at least some of the chlorine (to HCl).The hydrogenation of diolefins and olefins thus prevents, or at least limits, the formation of "gums," i.e., the polymerization of diolefins and olefins and therefore the formation of oligomers and polymers, which can clog the reaction section of hydrotreating step d). In parallel with hydrogenation, hydrometallation, and in particular the retention of silicon during step a), helps to limit the catalytic deactivation of the reaction section of hydrotreating step d). Furthermore, the conditions of step a) allow for the conversion of at least some of the chlorine, and preferably all of it.

[0100] A person skilled in the art will easily understand that, in step a) of hydrogenation, hydrogenation reactions such as those described above are carried out, but also in parallel some of the other hydrotreating reactions, and in particular hydrodesulfurization and hydrodeazotation, even if these reactions are rather favoured in step d) of hydrotreating, which is generally carried out at a higher temperature.

[0101] The temperature in step a), whether it be the average temperature (WABT), the temperature at the inlet of the reaction section or the temperature rise in step a) between the inlet and outlet of the reaction section, can in particular be controlled by injecting a diluent into step a), preferably a recycle of part of the liquid effluent from step b) and / or at least part of one or more cuts from step e), in particular by the recycle rate and / or by the temperature of the recycled effluent.

[0102] The temperature difference between the inlet and outlet of the reaction section of step a) includes the injection of a gaseous (hydrogen) or liquid cooling stream, preferably including a portion of the liquid effluent from step b).

[0103] The temperature difference between the inlet and outlet of the reaction section of step a) is exclusively due to the exothermicity of the chemical reactions carried out in the reaction section and is therefore understood to be without the use of a heating means (oven, heat exchanger etc).

[0104] Said reaction section implements hydrogenation in the presence of at least one hydrogenation catalyst, advantageously at an average temperature (or WABT as defined below) between 140 and 400°C, preferably between 240 and 350°C, and particularly preferably between 260 and 330°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs, preferably between 1.5 and 8.0 MPa abs. and at a volumetric rate per hour (WH) between 0.1 and 10.0 h1, preferably between 0.2 and 5.0 h1, and most preferably between 0.3 and 3.0 h1.

[0105] According to the invention, the "average temperature" of a reaction section corresponds to the Weight Average Bed Temperature (WABT), a term well known to those skilled in the art. The average temperature is advantageously determined based on the catalytic systems, equipment, and their configuration used. The average temperature (or WABT) is calculated as follows:

[0106] [Math.l] WABT = 1^X2

[0107] with Tinlet: the temperature of the flow at the inlet of the reaction section and Toutlet: the temperature of the effluent at the outlet of the reaction section. Unless otherwise specified, the "average temperature" of a reaction section is given at the start of the cycle.

[0108] The hourly volumetric velocity (WH) is defined here as the ratio between the hourly volumetric flow rate of the charge including the pyrolysis oil, possibly pre-treated, by the volume of catalyst(s).

[0109] Hydrogen coverage is defined as the ratio of the volumetric flow rate of hydrogen taken under normal temperature and pressure conditions to the volumetric flow rate of "fresh" feed, i.e. the feed to be treated, possibly pre-treated, without taking into account a recycled fraction, and in particular without taking into account the liquid effluent from step b) recycled, at 15°C (in normal m3, noted Nm3, of H2 per m3 of feed).

[0110] The quantity of the gaseous flow comprising hydrogen (H2), feeding said reaction section of step a), is advantageously such that the hydrogen coverage is between 100 and 1500 Nm3 of hydrogen per m3 of feed (Nm3 / m3), of preferably between 200 and 1000 Nm3 of hydrogen per m3 of charge (Nm3 / m3), preferably between 250 and 800 Nm3 of hydrogen per m3 of charge (Nm3 / m3).

[0111] Advantageously, the reaction section of said step a) comprises between 1 and 5 reactors, preferably between 2 and 5 reactors, and particularly preferably comprises two reactors. The advantage of a hydrogenation reaction section comprising several reactors lies in optimized feedstock processing, while reducing the risk of fouling of the catalytic bed(s) and thus preventing unit shutdown due to fouling.

[0112] According to a preferred embodiment, these reactors operate in a switchable mode, known as a "PRS" (Permutable Reactor System) or "lead and lag." Combining at least two reactors in PRS mode allows one reactor to be isolated, the spent catalyst to be discharged, the reactor to be refilled with fresh catalyst, and the reactor to be restarted without interrupting the process. The PRS technology is described, in particular, in patent FR2681871.

[0113] According to a particularly preferred embodiment, the hydrogenation reaction section of step a) comprises two reactors operating in switchable mode.

[0114] Advantageously, reactor internals, for example of the filter plate type, can be used to prevent clogging of the reactor(s). An example of a filter plate is described in patent FR3051375.

[0115] Advantageously, said hydrogenation catalyst comprises a support, preferably mineral, and a hydro-dehydrogenating function.

[0116] According to one embodiment, the hydro-dehydrogenating function comprises, in particular, at least one element from Group VIII, preferably selected from nickel and cobalt, and at least one element from Group VIB, preferably selected from molybdenum and tungsten. According to this embodiment, the total content, expressed as oxides of the metallic elements from Groups VIB and VIII, is preferably between 1% and 40% by weight, 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.

[0117] The weight ratio expressed in metal oxide between the metal (or metals) of group VIB and the metal (or metals) of group VIII is preferably between 1 and 20, and preferably between 2 and 10.

[0118] According to this embodiment, the reaction section of said step a) comprises, for example, a hydrogenation catalyst comprising between 0.5% and 12% by weight of nickel, preferably between 0.9% and 10% by weight of nickel (expressed as nickel oxide NiO relative to the weight of said catalyst), and between 1% and 30% by weight of molybdenum, Preference between 3% and 20% by weight of molybdenum (expressed as molybdenum oxide MoO3 relative to the weight of said catalyst) on a support preferably mineral, preferably on an alumina support.

[0119] According to another embodiment, the hydro-dehydrogenating function comprises, and preferably consists of, at least one element from Group VIII, preferably nickel. According to this embodiment, the nickel oxide content is preferably between 1 and 50% by weight, preferably between 10% and 30% by weight relative to the weight of said catalyst. This type of catalyst is preferably used in its reduced form, on a support, preferably mineral, preferably on an alumina support.

[0120] The support for said hydrogenation catalyst is preferably selected from alumina, silica, silica-aluminas, magnesia, clays, and mixtures thereof. This support may contain dopant compounds, in particular oxides selected from boron oxide, especially boron trioxide, zirconia, cerium, titanium dioxide, phosphoric anhydride, and mixtures thereof. Preferably, said hydrogenation catalyst comprises an alumina support, optionally doped with phosphorus and optionally with boron. When phosphoric anhydride (P₂O₅) is present, its concentration is less than 10% by weight relative to the weight of the alumina and advantageously at least 0.001% by weight relative to the total weight of the 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% relative to the total weight of alumina.The alumina used can be, for example, a y (gamma) or q (eta) alumina.

[0121] Said hydrogenation catalyst is for example in the form of extrudates.

[0122] Preferably, step a) may implement in addition to the ca hydrogenation catalysts described above, plus at least one hydrogenation catalyst used in step a) comprising less than 1% by weight of nickel and at least 0.1% by weight of nickel, preferably 0.5% by weight of nickel, expressed as nickel oxide (NiO) by weight of said catalyst, and less than 5% by weight of molybdenum and at least 0.1% by weight of molybdenum, preferably 0.5% by weight of molybdenum, expressed as molybdenum oxide (MoO3) by weight of said catalyst, on an alumina support. This low-metal content catalyst may preferably be placed upstream or downstream of the hydrogenation catalyst(s) described above, preferably upstream.

[0123] Preferably, step a) may implement, upstream of the hydrogenation catalyst(s), at least one guard bed containing adsorbents such as alumina, silica-alumina, zeolite, and / or activated carbon, possibly containing metals from groups VIB and / or VIII. Alternatively, a series of guard beds with particles of different diameters may be used, in particular a series of guard beds having decreasing diameters in the direction of charge flow (also called "grading" according to Anglo-Saxon terminology).

[0124] According to one embodiment, the feed, before preferably being mixed with at least a portion of the effluent from step b), can be preheated by direct heating to a temperature of up to 200°C, preferably up to 180°C, and particularly preferably up to 150°C. Above this temperature, contact with a wall during direct heating can induce the formation of gums and / or coke, which can cause fouling and an increase in the pressure drop of the feed heating system as well as of the catalyst bed(s). Heating the feed to a temperature above 150°C, preferably above 180°C, and particularly preferably above 200°C, is preferably carried out by indirect heating with a hot diluent, preferably with at least a portion of the effluent from step b).Thus, the temperature rise above 150°C, preferably above 180°C, and particularly preferably above 200°C, of ​​the feedstock is caused by mixing with a hotter liquid, and not by contact with a heated wall. This allows for localized limitation of high temperatures. This type of heating by mixing with a hot inert liquid therefore limits undesirable reactions such as the polymerization of diolefins (gum formation) and / or coke formation, and allows for adjusting the inlet temperature of the flow in step a) so as to initiate the hydrogenation reaction of unsaturates, preferably at the lowest possible temperature, while controlling the exothermicity of these reactions through a dilution effect on the reactive species.

[0125] According to another embodiment, the feed is entirely heated by indirect heating from at least a portion of the effluent from step b). In this case, the feed is not preheated before being mixed with at least a portion of the effluent from step b).

[0126] The energy required for the reaction and more specifically, the adjustment of the minimum temperature required for the activation of the saturation reactions of the double bonds is therefore advantageously achieved by mixing, upstream of step a), said feed comprising a pyrolysis oil and a hydrogen-rich gas, with a diluent and preferably a part of the liquid effluent from step b) of separation, having possibly undergone a temperature adjustment and preferably having been either preheated or cooled and particularly preferably having been preheated.

[0127] Another heating stream advantageously consists of a hydrogen-rich gaseous effluent from the hydrogen make-up and / or the gaseous effluent from separation step c). At least a portion of this hydrogen-rich gaseous effluent from the hydrogen make-up and / or the gaseous effluent from separation step c) is advantageously injected, optionally mixed with at least one part of the liquid effluent from step b) or separately, upstream of step a).

[0128] Said hydrogenation step a) yields a hydrogenated effluent, that is, an effluent with a reduced content of olefins, in particular diolefins, and metals, in particular silicon, and halogens, in particular chlorine. Hydrogenation step a) generally converts at least 40%, and preferably at least 60%, of the diolefins and at least 40%, and preferably at least 60%, of the olefins contained in the initial feed. The heat released by the saturation of the double bonds raises the temperature of the reaction medium and initiates the hydrotreating reactions, in particular the removal, at least in part, of other contaminants, such as silicon and chlorine or nitrogen. Preferably, at least 50%, and more preferably at least 75%, of the chlorine and silicon in the initial feed are respectively removed during step a). Generally, the silicon content is less than 10 ppm by weight.The effluent obtained at the end of step a) hydrogenation is sent, preferably directly, to step b) hot separation. (b) Hot separation stage

[0129] According to the invention, the treatment process comprises a separation step b), fed by the effluent from step a), said step being carried out at a temperature above the precipitation temperature of ammonium halides and at a pressure substantially identical to the pressure of step a) to obtain at least a first gaseous effluent, and a liquid effluent, part of which is preferably recycled upstream of step a).

[0130] The high temperature of this separation step b) prevents the precipitation of ammonium halide salts which are formed by reaction between halide ions and ammonium ions in order to recover a gaseous effluent containing the majority of the halides in the form of hydrogen halides (HCl) and gaseous ammonia, and a liquid effluent containing very little halides and ammonia.

[0131] The "precipitation temperature" of ammonium halides is defined as the temperature (under given conditions, such as concentration and pressure) at which gaseous ammonia and hydrogen halides precipitate, either by reacting to form solid ammonium halide crystals or by dissolving in water. The precipitation temperature depends on the halide concentrations and pressure according to thermodynamic principles. The precipitation temperature of ammonium halides is generally between 150 and 300°C, often between 180 and 295°C, and particularly between 200 and 290°C, under the conditions of use of the present process.

[0132] The temperature at which the separation in step b) is carried out must be higher than the precipitation temperature of ammonium halides in order to recover a The gaseous effluent contains the majority of the halides in the form of hydrogen halides (HCl) and gaseous ammonia, and the liquid effluent contains very little halide and ammonia. The temperature at which the separation in step b) is carried out is generally between 200 and 450°C, preferably between 220 and 330°C, and particularly preferably between 240 and 300°C.

[0133] The term "pressure substantially equal to the pressure of step a)" means the pressure of step a) with a pressure difference of between 0 and 1 MPa, preferably between 0.005 and 0.3 MPa, and particularly preferably between 0.01 and 0.3 MPa with respect to the pressure of step a). Preferably, the pressure of step b) is the pressure of step a) less the pressure losses.

[0134] Step b) of separation can advantageously be carried out by any method known to those skilled in the art, such as, for example, the combination of one or more separator(s) (balloon(s)), and / or one or more stripping column(s), this separator(s) (balloon(s)) and / or column(s) optionally being supplied with a stripping gas, for example, a hydrogen-rich gas stream. Preferably, step b) is carried out with a single separator (balloon).

[0135] The separation step b) is a so-called high-pressure or medium-pressure, high-temperature separation step, also known to those skilled in the art as HHPS (for "Hot High Pressure Separator" according to Anglo-Saxon terminology). Thus, this step b) preferably uses a so-called "hot high-pressure" separator, the pressure being substantially equal to the operating pressure of step a).

[0136] The gas / liquid separation has an efficiency corresponding to the solubility and Henry's law. This means that an equilibrium amount of halides in the form of hydrogen halides will remain in the liquid effluent obtained in step b). These halides will be released into the gaseous effluent during step c) of cold separation and then removed by dissolution in the aqueous effluent.

[0137] Advantageously, part of the liquid effluent from step b) is recycled upstream of step a).

[0138] Mixing the feedstock with a portion of the liquid effluent (hot recycle) upstream of step a) allows, on the one hand, for the dilution of impurities in the feedstock and, on the other hand, for temperature control in step a), where highly exothermic reactions (in particular, the hydrogenation reactions of olefins and diolefins) occur. Mixing the feedstock with a portion of the liquid effluent (hot recycle) upstream of step a) also allows for indirect heating of the feedstock, simply by mixing the "cold" feedstock with the hot recycle.

[0139] The quantity of recycled liquid effluent from step b) is adjusted so that the weight ratio between the recycled stream from step b) and the feed including pyrolysis oil, i.e. the feed to be treated supplying the overall 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 quantity of recycled liquid effluent from step b) is adjusted so that the weight ratio between the recycle stream and the feed including pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5. This recycle rate allows control of the temperature rise in step a). Indeed, when the recycle rate is high, the feed dilution rate is high, and the temperature rise at the beginning of the reaction section of step a), particularly due to the hydrogenation reactions of diolefins, is thus controllable by the dilution effect.

[0140] According to a preferred embodiment, at least a portion of the liquid effluent from step b) may advantageously be either cooled, or preheated, if necessary, or kept at the same temperature as at the outlet of separation step b), before preferably being recycled upstream of hydrogenation step a), according to the temperature and flow rate of feed and hydrogen, so that the temperature of the incoming stream, comprising said feed preferably mixed with at least a portion of said liquid effluent from step b) and a hydrogen-rich gas, is between 140 and 400°C, preferably between 220 and 350°C, and particularly preferably between 260 and 330°C.

[0141] High-pressure, high-temperature separation allows, on the one hand, for maximizing energy recovery by hot recycling of a portion of the liquid effluent. Indeed, the energy required to reach the inlet temperature in step a) is at least partially supplied by the heat from a portion of the liquid effluent from step b) and also reduces or even eliminates the need for preheating by directly heating the feed above 200°C to prevent gum formation. Furthermore, preferably recycling at least a portion of the high-pressure liquid effluent saves energy for its pressurization in step a).

[0142] High-pressure, high-temperature separation also minimizes the amount of light fraction (naphtha cut) contained in the liquid effluent. At this temperature, almost all of the light fraction of the effluent (naphtha) is released as a gaseous effluent to step c) of cold separation / washing, while the liquid phase consists mainly of the heavy fraction of the feedstock (middle distillates cut). When recycling part of the liquid effluent to step a), the ppH2 is favored in step a) because the light fraction (naphtha) could partially vaporize and lower the ppH2 if it were not at least partially removed during the high-pressure, high-temperature separation. The removal of the light fraction, including naphtha, This can optionally be increased by a slight expansion upstream of at least one separator implemented in step b), although this implementation is not preferred due to the energy loss associated with the expansion. Another option for increasing the removal of the light fraction including naphtha could be to perform stripping, for example by injecting a hydrogen-rich gas in step b). (c) Cold separation and washing stage

[0143] According to the invention, the treatment process comprises a separation step c), fed by the first gaseous effluent and optionally at least a portion of the liquid effluent from step b) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of ammonium halides, and at a pressure substantially the same as or lower than the pressure of step b), to obtain at least a second gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

[0144] Separation in step c) at a temperature lower than the precipitation temperature of ammonium halides from the gaseous effluent mixture and possibly at least part of the liquid effluent from step b) causes the ammonium halide salts which are formed by reaction between halide ions and ammonium ions contained in particular in the gaseous effluent.

[0145] Cold separation step c) is fed with the gaseous effluent alone from step b) or with the mixture of the gaseous effluent and at least a portion of the liquid effluent from step b). Preferably, cold separation step c) is fed with the mixture of the gaseous effluent and at least a portion of the liquid effluent. According to a particularly preferred embodiment, cold separation step c) is fed with the mixture of the gaseous effluent and all of the liquid effluent. Performing cold separation step c) on the mixture of gaseous and liquid effluent has the advantage of extracting not only contaminants from the gaseous effluent, but also hydrogen halides (HCl) dissolved in the liquid effluent. Indeed, the hydrogen halides dissolved in the liquid effluent from step b) are released and also form ammonium halides with the ammonia present in the gaseous effluent.

[0146] Washing this step c) with an aqueous solution 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 in which the ammonium halide salts are dissolved.

[0147] The temperature at which the separation in step c) is carried out must be lower than the precipitation temperature of ammonium halides in order to precipitate the ammonium halide salts. The temperature at which the separation in step c) is carried out is greater than or equal to 20°C and less than 200°C, preferably between 25 and 120°C, and particularly preferably between 30 and 70°C.

[0148] "Pressure substantially equal to the pressure of step b)" means the pressure of step b) with a pressure difference of between 0 and 1 MPa, preferably between 0.005 and 0.3 MPa, and particularly preferably between 0.01 and 0.3 MPa, relative to the pressure of step b). Preferably, the pressure of step c) is the pressure of step b) less the pressure losses. Furthermore, operating at least part of the separation step c) at a pressure substantially identical to the operating pressure of step b) facilitates the recycling of hydrogen contained in the gaseous effluent.

[0149] The separation step c) can also be carried out at a pressure lower than the pressure of step b).

[0150] The separation step c) 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 carried out at the same or lower temperature and pressure as each separation step of the preceding step c).

[0151] Step c) of separation can advantageously be carried out by any method known to those skilled in the art, such as, for example, the combination of one or more separator(s) (balloon(s)), and / or one or more stripping column(s), this separator(s) (balloon(s)) and / or column(s) optionally being supplied with a stripping gas, for example, a hydrogen-rich gas stream. Preferably, step c) is carried out with a single separator (balloon).

[0152] The separation step c) is preferably carried out in at least one so-called high-pressure or medium-pressure, low-temperature separator vessel, also known to those skilled in the art as CHPS (for "Cold High Pressure Separator" according to Anglo-Saxon terminology). Thus, this step c) preferably uses a so-called "cold high-pressure" separator, the pressure being substantially equal to the operating pressure of step b).

[0153] The washing / separation section of step c) can be at least partially carried out in common or separate washing and separation equipment, such equipment being well known (separating vessels capable of operating at different pressures and temperatures, pumps, heat exchangers, scrubbing columns, etc.). The separation step c) may, for example, include a column for stripping acid water (also called a "sour water stripper" in Anglo-Saxon terminology) from the withdrawn aqueous effluent, a column for scrubbing acid gases to purify the hydrogen-rich gas before recycling, and a column for stabilizing the washed liquid effluent to remove dissolved gases.

[0154] The gaseous effluent obtained at the end of step c) advantageously comprises The hydrogen, preferably comprising at least 80% by volume, preferably at least 85% by volume, of hydrogen. Advantageously, said gaseous effluent may at least partially be recycled to the steps a) hydrogenation and / or d) hydrotreating and / or to one or both steps f) hydrocracking when present, the recycling system being able to include a purification section.

[0155] The gaseous effluent can also be subjected to further separation(s) in order to recover at least one hydrogen-rich gas and / or light hydrocarbons, in particular ethane, propane and butane, which can advantageously be sent separately or in a mixture to one or more furnaces of step g) of steam cracking so as to increase the overall olefin yield.

[0156] The aqueous effluent obtained at the end of step c) advantageously comprises dissolved ammonium salts and / or hydrogen halides (for example hydrochloric acid).

[0157] Depending on the halogen content of the initial feed to be treated, a stream containing an amine such as, for example, monoethanolamine, diethanamine, and / or monodiethanolamine can be injected upstream of step a) of hydrogenation to ensure a sufficient quantity of ammonium ions to combine the halide ions formed during the hydrogenation step, thereby limiting the formation of hydrochloric acid and thus limiting corrosion downstream of the separation section. This injection may be necessary, in particular, when the feed contains a large amount of halogenated compounds or few nitrogen compounds, or when the hydrogenation of the nitrogen compounds is not sufficiently advanced during step a) to form enough gaseous ammonia or ammonium ions.

[0158] In a possible embodiment of the invention, step c) of separation comprises the injection of an aqueous solution into the mixture of the gaseous effluent and another portion of the liquid effluent from step b), followed by the washing / separation section advantageously comprising a separation phase enabling the production of at least one aqueous effluent containing dissolved ammonium halide salts, one washed liquid hydrocarbon effluent, and one partially washed gaseous effluent. The aqueous effluent containing dissolved ammonium halide salts and the washed liquid hydrocarbon effluent can then be separated in a settling tank to obtain said hydrocarbon effluent and said aqueous effluent.The said partially washed gaseous effluent can simultaneously be introduced into a scrubbing column where it flows counter-currently to an aqueous flow, preferably of the same nature as the aqueous solution injected into the hydrocarbon effluent, thus allowing the removal, at least in part, and preferably in full, of hydrogen halides (of the HCl type) contained in the partially washed gaseous effluent and thus obtaining the said gaseous effluent, preferably comprising essentially hydrogen, and an acidic aqueous flow. The said effluent. The aqueous solution from the settling tank may optionally be mixed with said acidic aqueous stream and used, possibly mixed with said acidic aqueous stream, in a water recycling circuit to supply step c) of the separation of said aqueous solution upstream of the washing / separation section and / or said aqueous stream in the washing column. Said water recycling circuit may include a water top-up and / or a basic solution and / or a purge to remove dissolved salts.

[0159] The hydrocarbon effluent from the separation step c) is sent, in part or in whole, to the hydrotreatment step d). (d) Hydrotreatment stage

[0160] According to the invention, the treatment process comprises a hydrotreating step d) carried out in a hydrotreating reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrotreating catalyst, said hydrotreating reaction section being fed at least by said hydrocarbon effluent from step c), optionally by at least a part of the liquid effluent from step b), and a second gaseous stream comprising hydrogen, said hydrotreating reaction section being carried out at an average temperature between 250 and 430°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a hydrotreated effluent.

[0161] Advantageously, step d) implements hydrotreating reactions well known to those skilled in the art, and more particularly hydrotreating reactions such as aromatic hydrogenation, hydrodesulfurization, and hydrodeazotation. Furthermore, the hydrogenation of the remaining olefins and halogenated compounds, as well as hydrodemetallation, can proceed even if most, and preferably all, of these impurities has been removed in step a).

[0162] Said hydrotreating reaction section is advantageously implemented at a pressure equivalent to that used in the reaction section of step a) of hydrogenation, and generally at a higher average temperature than that of the reaction section of step a) of hydrogenation. Thus, said hydrotreating reaction section is advantageously implemented at an average hydrotreating temperature between 250 and 430°C, preferably between 280 and 380°C, at a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and at a volumetric flow rate per hour (WH) between 0.1 and 10.0 h⁻¹, preferably between 0.1 and 5.0 h⁻¹, preferably between 0.2 and 2.0 h⁻¹, most preferably between 0.2 and 1 h⁻¹. The hydrogen coverage in step d) is advantageously between 100 and 1500 Nm³ of hydrogen per m³ of fresh feedstock supplying step a), and preferably between 200 and 1000 Nm3 of hydrogen per m3 of fresh feed supplying stage a), preferably between 250 and 800 Nm3 of hydrogen per m3 of fresh feed supplying stage a). The definitions of mean temperature (WABT), WH and hydrogen cover correspond to those described above.

[0163] Said hydrotreating reaction section is fed at least by said hydrocarbon effluent from step c) and a second gaseous stream comprising hydrogen, advantageously at the level of the first catalytic bed of the first operating reactor.

[0164] Optionally, the reaction section of said step d) may also be further supplied by at least part of the liquid effluent from step b). This has the advantage of saving energy because the liquid effluent from step b) is hot and does not need to be heated before being introduced into step d).

[0165] According to another (preferred) variant and where step c) of cold separation is fed by the mixture of gaseous effluent and at least part, and preferably all, of the liquid effluent from step b), the reaction section of said step d) is fed only by said hydrocarbon effluent from step c).

[0166] Advantageously, said step d) is carried out in a hydrotreating reaction section comprising at least one, preferably between one and five, fixed-bed reactor(s) having n catalytic beds, n being an integer greater than or equal to one, preferably between one and ten, preferably between two and five, said bed(s) each comprising at least one, and preferably not more than ten, hydrotreating catalysts. When a reactor comprises several catalytic beds, that is to say at least two, preferably between two and ten, preferably between two and five catalytic beds, said catalytic beds are preferably arranged in series in said reactor.

[0167] When step d) is implemented in a hydrotreating reaction section comprising several, preferably two, reactors, these reactors can operate in series and / or in parallel and / or in switchable (or PRS) mode and / or in swing mode. The various possible operating modes, PRS (or lead and lag) mode and swing mode, are well known to those skilled in the art and are advantageously defined above.

[0168] In another embodiment of the invention, said hydrotreating reaction section comprises a single fixed-bed reactor containing n catalytic beds, n being an integer greater than or equal to one, preferably between one and ten, preferably between two and five.

[0169] In a particularly preferred mode, the hydrogenation reaction section of step a) comprises two reactors operating in switchable mode followed by the hydrotreating reaction section of step d) which comprises a single fixed-bed reactor.

[0170] Advantageously, said hydrotreating catalyst used in said step d) can be selected from known hydrodemetallation, hydrotreating, silicon capture catalysts, used in particular for the treatment of petroleum cuts, and combinations thereof. Examples of known hydrometallation catalysts are those described in patents EP 0113297, EP 0113284, US 5221656, US 5827421, US 7119045, US 5622616, and US 5089463. Examples of known hydrotreating catalysts are those described in patents EP 0113297, EP 0113284, US 6589908, US 4818743, and US 6332976. Examples of known silicon capture catalysts are those described in patent applications CN 102051202 and US 2007 / 080099.

[0171] In particular, said hydrotreating catalyst comprises a support, preferably mineral, and at least one metallic element having a hydro-dehydrogenating function. Said metallic element having a hydro-dehydrogenating function advantageously comprises at least one element from Group VIII, preferably selected from the group consisting of nickel and cobalt, and / or at least one element from Group VIB, preferably selected from the group consisting of molybdenum and tungsten. The total content, expressed as oxides of the metallic elements from Groups VIB and VIII, is preferably between 0.1% and 40% by weight, 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, expressed as metal oxide, of the metal(s) of Group VIB to the metal(s) of Group VIII is preferably between 1.0 and 20, preferably between 2.0 and 10. For example, the hydrotreating reaction section of step d) of the process comprises a hydrotreating catalyst comprising between 0.5% and 10% by weight of nickel, preferably between 1% and 8% by weight of nickel, expressed as nickel oxide NiO relative to the total weight of the hydrotreating catalyst, and between 1.0% and 30% by weight of molybdenum, preferably between 3.0% and 29% by weight of molybdenum, expressed as molybdenum oxide MoO3 relative to the total weight of the hydrotreating catalyst, on a mineral support, preferably on an alumina support.

[0172] The support for said hydrotreating catalyst is advantageously selected from alumina, silica, silica-aluminas, magnesia, clays, and mixtures thereof. This support may further contain doping compounds, in particular oxides selected from boron oxide, especially boron trioxide, zirconia, cerium, titanium dioxide, phosphoric anhydride, 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 alumina and advantageously at least 0.001% by weight relative to the total alumina. When boron trioxide (B₂O₃) is present, its concentration is less than 10% by weight relative to the alumina and advantageously at least 0.001% by weight relative to the total alumina. The alumina used may be, for example, gamma (γ) or eth (β) alumina.

[0173] Said hydrotreating catalyst is for example in the form of extrudates.

[0174] Advantageously, said hydrotreating catalyst of the process has a specific surface area greater than or equal to 250 m² / g, preferably greater than or equal to 300 m² / g. The specific surface area of ​​said hydrotreating catalyst is advantageously less than or equal to 800 m² / g, preferably less than or equal to 600 m² / g, and in particular less than or equal to 400 m² / g. The specific surface area of ​​the hydrotreating catalyst is measured by the BET method, that is, the specific surface area determined by nitrogen adsorption in accordance with ASTM D 3663-78, established from the Brunauer-Emmett-Teller method described in the periodical 'The Journal of the American Chemical Society', 6Q, 309 (1938). Such a specific surface area further improves the removal of contaminants, in particular metals such as silicon.

[0175] According to another aspect of the invention, the hydrotreating catalyst as described above further comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur. Such a catalyst is often referred to as an "additized catalyst." Generally, the organic compound is chosen from a compound having one or more chemical functions selected from among a carboxyl group, alcohol, thiol, thioether, sulfone, sulfoxide, ether, aldehyde, ketone, ester, carbonate, amine, nitrile, imide, oxime, urea, and amide, or compounds including a furan ring, or sugars.

[0176] Preferably, step d) may implement upstream of the hydrogenation catalyst(s) at least one guard bed or a series of guard beds of the "grading" type as described above for step a).

[0177] Advantageously, the hydrotreating step (d) enables the hydrogenation of at least 80%, and preferably all, of the olefins and halogenated compounds remaining after the hydrogenation step (a), but also the conversion, at least in part, of other impurities present in the feedstock, such as aromatic compounds, metallic compounds, sulfur compounds, nitrogen compounds, and oxygenated compounds. Preferably, the nitrogen content at the outlet of step (d) is less than 100 ppm by weight, and Preferably less than 10 ppm by weight. Preferably, the sulfur content at the outlet of step d) is less than 100 ppm by weight, and preferably less than 10 ppm by weight. Step d) can also further reduce the contaminant content, such as that of metals, particularly silicon. Preferably, the metal content at the outlet of step d) is less than 10 ppm by weight, and 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 d) is less than 5 ppm by weight.

[0178] Depending on the sulfur content of the initial feedstock to be treated, a stream containing a sulfurizing agent may be injected upstream of step a) hydrogenation and / or step d) hydrotreating and / or upstream of one of the hydrocracking steps when present, preferably upstream of step a) hydrogenation and / or step d) hydrotreating in order to ensure a sufficient quantity of sulfur to form or maintain the active species of the catalyst (in sulfur form). This activation or sulfurization step is carried out by methods well known to those skilled in the art, and advantageously under a sulfur-reducing atmosphere in the presence of hydrogen and hydrogen sulfide.The sulfurizing agents are preferably hydrogen sulfide (H₂S), elemental sulfur, CS₂, mercaptans, sulfides and / or polysulfides, hydrocarbon fractions with a boiling point below 400°C containing sulfur compounds, or any other sulfur-containing compound used for activating hydrocarbon feedstocks to sulfide the catalyst. These sulfur-containing compounds are advantageously selected from alkyl disulfides, such as dimethyl disulfide (DMDS), alkyl sulfides, such as dimethyl sulfide, thiols, such as n-butylmercaptan (or 1-butanethiol), and polysulfide compounds of the tertiononyl 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 a sulfurizing agent and a hydrocarbon feedstock.Preferably, the catalyst is sulfided in situ in the presence of the added charge of dimethyl disulfide.

[0179] The hydrotreated effluent from step d) of hydrotreatment can be sent, in part or in whole, either to a step e) of fractionation, or directly to a step f) of hydrocracking, or directly to a step g) of steam cracking, or directly into a fuel storage unit.

[0180] Sending the hydrotreated effluent from step d) of hydrotreatment directly into a hydrocracking step has the advantage of not having to repressurize it.

[0181] According to another variant, the hydrotreated effluent can undergo a hot separation step followed by a cold separation step with washing in the same way as the steps b) and c) described above. This variant is advantageous for feedstocks with high concentrations of halogenated compounds, particularly chlorinated compounds. The hydrocarbon effluent exiting the cold separation with washing can then be sent, in part or in whole, either to a fractionation step e), or directly to a hydrocracking step f), or directly to a steam cracking step g), or directly to a fuel storage unit.

[0182] Advantageously, a pump can be used between the separation step (c) and the hydrotreating step (d) so as to raise the total pressure at the inlet of (d). Advantageously, an additional flow of hydrogen can be injected at the inlet of the hydrotreating step (d) so as to adjust the partial pressure of hydrogen. (e) Fractionation step (optional)

[0183] The process according to the invention may include a fractionation step of all or part, preferably all, of the hydrotreated effluent from step d), to obtain at least a third gas stream, a naphtha cut and at least a middle distillates cut.

[0184] The term "naphtha cut" means a hydrocarbon cut comprising compounds having a boiling point generally less than or equal to 175°C, in particular between 80 and 175°C.

[0185] The term "middle distillate cut" means a hydrocarbon cut comprising compounds having a boiling point generally above 175°C. The heavy cut may include middle distillates such as a diesel cut and / or a kerosene cut. It may also include heavier compounds.

[0186] Depending on the destination or use of the fractions from step e) of fractionation, a person skilled in the art will adjust the cut points in the stripping and / or distillation operations. For example, it may be necessary to adjust the endpoint of the naphtha cut to 150, 175 or 200°C.

[0187] Step e) in particular allows the removal of gases dissolved in the hydrotreated liquid effluent, such as ammonia, hydrogen sulfide and light hydrocarbons having 1 to 4 carbon atoms.

[0188] The optional fractionation step e) is advantageously carried out at a pressure less than or equal to 1.0 MPa abs., preferably between 0.1 and 1.0 MPa abs.

[0189] According to one embodiment, step e) can be carried out in a section advantageously comprising at least one stripping column equipped with a reflux circuit including a reflux flask. Said stripping column is fed by the hydrotreated liquid effluent from step d) and by a stream of steam. The hydrotreated liquid effluent from step d) may optionally be heated before entering the stripping column. Thus, the lighter compounds are carried to the top of the column and into the reflux circuit including a reflux flask. in which a gas / liquid separation takes place. The gaseous phase, which includes the light hydrocarbons, is withdrawn from the reflux flask as a gas stream. The naphtha fraction is advantageously withdrawn from the reflux flask as a liquid stream. The middle distillates fraction is advantageously withdrawn from the bottom of the stripping column.

[0190] According to another embodiment, step e) can be carried out in a section comprising one or more separator tank(s). Advantageously, at least one low-pressure, low-temperature separator tank will be used to remove gases dissolved in the hydrotreated liquid effluent.

[0191] According to other embodiments, step e) of fractionation may employ a stripping column or a balloon separator followed by a distillation column or only a distillation column.

[0192] The naphtha cut and the middle distillates cut, possibly mixed, may be sent, in whole or in part, to a steam cracking unit, at the end of which olefins may be (re)formed to participate in polymer formation. Preferably, only a portion of said cuts is sent to a steam cracking unit; at least a fraction of the remaining portion is optionally recycled in at least one of the process steps and / or sent to a fuel storage unit, for example a naphtha storage unit, a diesel storage unit, or a kerosene storage unit, derived from conventional petroleum feedstocks.

[0193] According to a preferred method, the naphtha cut, in whole or in part, is sent to a steam cracking unit, while the middle distillates cut is sent to a hydrocracking step f) and / or sent to a fuel storage unit.

[0194] In a particular embodiment, the optional fractionation step e) can provide, in addition to a gas stream, a naphtha cut (generally comprising compounds having a boiling point less than or equal to 175°C, preferably between 80 and 175°C), and a middle distillates cut (generally comprising compounds having a boiling point greater than 175°C and less than 385°C), and a heavy hydrocarbon cut (generally comprising compounds having a boiling point greater than or equal to 385°C).The naphtha cut can be sent, in whole or in part, to a steam cracking unit and / or to the naphtha storage unit from conventional petroleum feedstocks, or it can also be recycled; the middle distillate cut can also be sent, in whole or in part, either to a steam cracking unit, or to a diesel storage unit from conventional petroleum feedstocks, or to the hydrocracking stage f) when present, or it can also be recycled; the heavy cut can, for its part, be sent, at least in part, to a steam cracking unit, or be sent to the hydrocracking stage when present.

[0195] In another particular embodiment, the optional step e) of frac The process can yield, in addition to a gas stream, a naphtha cut (generally comprising compounds with a boiling point of 175°C or lower, preferably between 80 and 175°C), a kerosene cut (generally comprising compounds with a boiling point above 175°C and below 280°C), a diesel cut (generally comprising compounds with a boiling point above 280°C and below 385°C), and a heavy hydrocarbon cut (generally comprising compounds with a boiling point above or equal to 385°C). The naphtha cut, kerosene cut, and / or diesel cut can be, in whole or in part, either sent to a steam cracking unit, or respectively to a naphtha, kerosene, or diesel pool derived from conventional petroleum feedstocks, or recycled. Diesel and / or kerosene cuts can also be sent to the hydrocracking stage f) when it is present.The heavy cut can, at least in part, be sent to a steam cracking unit, or be sent to the hydrocracking stage (f) when present.

[0196] In another particular embodiment, the naphtha cut from step e) is fractionated into a heavy naphtha cut (generally comprising compounds having a boiling point between 80 and 175°C) and a light naphtha cut (generally comprising compounds having a boiling point below 80°C), at least a portion of said heavy naphtha cut being sent to an aromatic complex comprising at least one naphtha reforming step for the production of aromatic compounds. According to this embodiment, at least a portion of the light naphtha cut is sent to the steam cracking step g) described below.

[0197] The gaseous effluent(s) from fractionation step e) may be subjected to further purification and separation in order to recover at least light hydrocarbons, in particular ethane, propane, and butane, which may advantageously be sent separately or mixed to one or more furnaces of steam cracking step g) so as to increase the overall olefin yield. (f) Hydrocracking step (optional)

[0198] According to one variant, the process of the invention may include a hydrocracking step f) carried out after hydrotreating step d) with at least a portion of said hydrotreated effluent from step d) or carried out after fractionating step e) with at least a portion of the middle distillate cut.

[0199] Advantageously, step f) implements hydrocracking reactions well known to those skilled in the art, and more particularly allows the conversion of heavy compounds, for example compounds having a boiling point above 175°C, into compounds having a boiling point less than or equal to 175°C. Other reactions, such as the hydrogenation of olefins, aromatics, hydrodemetallation, hydrodesulfurization, hydrodeazotation, etc., can be carried out.

[0200] Compounds with a boiling point above 175°C contain, compared to lighter compounds, more naphthenic, naphtheno-aromatic, and aromatic compounds, thus leading to a higher C / H ratio. This high ratio causes coking in the steam cracker, therefore requiring steam cracking furnaces dedicated to this cut. When it is desired to minimize the yield of these heavy compounds (middle distillate cut) and maximize the yield of light compounds (naphtha cut), these compounds can be converted, at least partially, into light compounds by hydrocracking, a cut generally preferred for a steam cracking unit.

[0201] Thus, the process of the invention may comprise a hydrocracking step (f) carried out in a hydrocracking reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrocracking catalyst, said hydrocracking reaction section being fed by at least a portion of said hydrotreated effluent from step (d) and / or by at least a portion of the middle distillate cut from step (e) and a third gas stream comprising hydrogen, said hydrocracking reaction section being carried out at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs., and an hourly volumetric rate between 0.1 and 10.0 h⁻¹ to obtain a first hydrocracking effluent.

[0202] Thus, said hydrocracking reaction section is advantageously implemented at an average temperature between 250 and 480°C, preferably between 320 and 450°C, at a partial pressure of hydrogen between 1.5 and 20.0 MPa abs., preferably between 3 and 18.0 MPa abs., and at a volumetric rate per hour (WH) between 0.1 and 10.0 h⁻¹, preferably between 0.1 and 5.0 h⁻¹, preferably between 0.2 and 4 h⁻¹. The hydrogen coverage in step f) is advantageously between 80 and 2000 Nm³ of hydrogen per m³ of fresh feedstock supplying step f), and preferably between 200 and 1800 Nm³ of hydrogen per m³ of fresh feedstock supplying step f). The definitions of mean temperature (WABT), WH and hydrogen cover correspond to those described above.

[0203] Advantageously, said hydrocracking reaction section is implemented at a pressure equivalent to that used in the reaction section of step a) of hydrogenation or step d) of hydrotreating.

[0204] Advantageously, said step f) is carried out in a hydrocracking reaction section comprising at least one, preferably between one and five, fixed-bed reactor(s) having n catalytic beds, n being an integer greater than or equal to one, preferably between one and ten, more preferably between two and five, said bed(s) each comprising at least one, and preferably not more than ten, hydrocracking catalysts. When a reactor comprises several catalytic beds, that is to say at least two, preferably between two and ten, preferably between two and five catalytic beds, said catalytic beds are preferably arranged in series in said reactor.

[0205] The hydrocracked effluent can at least partly be recycled in step a) hydrogenation and / or in step b) hot separation and / or in step c) hot separation and / or in step d) hydrotreating and / or in step e) fractionation.

[0206] The hydrocracking step can be carried out in one (step f) or two steps (steps f) and f'). When carried out in two steps, the effluent from the first hydrocracking step f) is separated to obtain a hydrocarbon fraction comprising compounds with a boiling point above 175°C (middle distillate fraction), which is introduced into the second hydrocracking step f', comprising a second dedicated hydrocracking reaction section, different from the first hydrocracking reaction section f). This configuration is particularly suitable when it is desired to produce only a naphtha fraction.

[0207] The second hydrocracking step f') is carried out in a hydrocracking reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrocracking catalyst. The hydrocracking reaction section is fed with at least a portion of the first hydrocracking effluent from the first hydrocracking step f) and a gas stream comprising hydrogen. The hydrocracking reaction section is operated at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs., and a volumetric flow rate between 0.1 and 10.0 h⁻¹, to obtain a second hydrocracking effluent. The preferred operating conditions and catalysts used in the second hydrocracking step are those described for the first hydrocracking step.The operating conditions and catalysts used in the two hydrocracking stages may be identical or different.

[0208] Said second hydrocracking step is preferably carried out in a hydrocracking reaction section comprising at least one, preferably between one and five, fixed bed reactor(s) having n catalytic beds, n being an integer greater than or equal to one, preferably between one and ten, preferably between two and five, said bed(s) each comprising at least one, and preferably not more than ten, hydrocracking catalyst(s).

[0209] These operating conditions used in the hydrocracking step(s) generally make it possible to achieve conversions per pass, into products having at least 80% by volume of compounds having boiling points less than or equal to 175°C, preferably below 160°C and preferably below 150°C, and above 15 wt% and even more preferably between 20 and 95 wt%. When the process is carried out in two hydrocracking stages, the conversion per pass in the second stage is kept moderate to maximize selectivity for naphtha cut compounds (having a boiling point below or equal to 175°C, in particular between 80 and below or equal to 175°C). The conversion per pass is limited by the use of a high recycle ratio on the loop of the second hydrocracking stage. This ratio is defined as the ratio between the feed flow rate of stage f') and the feed flow rate of stage a); preferably this ratio is between 0.2 and 4, preferably between 0.5 and 2.5.

[0210] The hydrocracked effluent from the second hydrocracking stage f') can at least partly be recycled in the hydrogenation stage a) and / or in the hot separation stage b) and / or in the hot separation stage c) and / or in the hydrotreating stage d) and / or in the fractionation stage e).

[0211] The hydrocracking step(s) do not necessarily allow the conversion of all hydrocarbon compounds in the middle distillates cut into hydrocarbon compounds with a boiling point of 175°C or lower (naphtha cut). After the fractionation step e), a more or less significant proportion of compounds with a boiling point above 175°C may therefore remain. At least part of this unconverted cut can be introduced into a second hydrocracking step f'). Another part can be purged. Depending on the process operating conditions, this purging can be between 0 and 10% by weight of the cut containing compounds with a boiling point above 175°C relative to the incoming feed, and preferably between 0.5% and 5% by weight.

[0212] According to the invention, the hydrocracking step(s) operate in the presence of at least one hydrocracking catalyst.

[0213] The hydrocracking catalyst(s) used in the hydrocracking step(s) are conventional hydrocracking catalysts known to those skilled in the art, of a bifunctional type combining an acid function with a hydro-dehydrogenating function and optionally at least one binding matrix. The acid function is provided by large surface area supports (generally 150 to 800 m² / g) exhibiting surface acidity, such as halogenated aluminas (particularly chlorinated or fluorinated), combinations of boron and aluminum oxides, amorphous silica-aluminas, and zeolites. The hydro-dehydrogenating function is provided by at least one metal from group VIB of the periodic table and / or at least one metal from group VIII.

[0214] Preferably, the hydrocracking catalyst(s) comprise a hydro-dehydrogenating function comprising at least one metal from Group VIII selected from the iron, cobalt, nickel, ruthenium, rhodium, palladium, and platinum, and preferably cobalt and nickel. Preferably, said catalysts also comprise at least one metal from group VIB selected from chromium, molybdenum, and tungsten, alone or in mixtures, and preferably from molybdenum and tungsten. Hydro-dehydrogenating functions of the NiMo, NiMoW, and NiW types are preferred.

[0215] Preferably, the content of Group VIII metal in the hydrocracking catalyst(s) is advantageously between 0.5 and 15 wt% and preferably between 1 and 10 wt%, the percentages being expressed as a wt% of oxides 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.

[0216] Preferably, the content of group VIB metal in the hydrocracking catalyst(s) is advantageously between 5 and 35 wt%, and preferably between 10 and 30 wt%, the percentages being expressed as a wt% of oxides 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.

[0217] The hydrocracking catalyst(s) may also optionally include at least one promoting element deposited on the catalyst and selected from the group formed by phosphorus, boron and silicon, optionally at least one element from group VIIA (chlorine, fluorine preferred), optionally at least one element from group VIIB (manganese preferred), and optionally at least one element from group VB (niobium preferred).

[0218] Preferably, the hydrocracking catalyst(s) comprise at least one amorphous or poorly crystallized porous mineral matrix of the oxide type selected from aluminas, silicas, silica-aluminas, aluminates, alumina-boron oxide, magnesia, silica-magnesia, zirconia, titanium oxide, clay, alone or in mixture, and preferably aluminas or silica-aluminas, alone or in mixture.

[0219] Preferably, the silica-alumina contains more than 50% by weight of alumina, preferably more than 60% by weight of alumina.

[0220] Preferably, the hydrocracking catalyst(s) may also include a zeolite selected from the Y zeolites, preferably from the US Y zeolites, alone or in combination with other zeolites from the beta zeolites, ZSM-12, IZM-2, ZSM-22, ZSM-23, SAPO-11, ZSM-48, ZBM-30, alone or in mixtures. Preferably, the zeolite is the USY zeolite alone.

[0221] In the case where said catalyst includes a zeolite, the zeolite content in the hydrocracking catalyst(s) is advantageously between 0.1 and 80% by weight, preferably between 3 and 70% by weight, the percentages being expressed as a percentage of zeolite relative to the total weight of the catalyst.

[0222] A preferred catalyst comprises, and preferably consists of, at least one metal from group VIB and optionally at least one non-noble metal from group VIII, at least one promoting element, preferably phosphorus, at least one zeolite Y and at least one alumina binder.

[0223] An even more preferred catalyst comprises, and is preferably made up of, nickel, molybdenum, phosphorus, a USY zeolite, and possibly also a beta zeolite, and alumina.

[0224] Another preferred catalyst comprises, and is preferably made up of, nickel, tungsten, alumina and silica-alumina.

[0225] Another preferred catalyst comprises, and is preferably made up of, nickel, tungsten, USY zeolite, alumina and silica-alumina.

[0226] Said hydrocracking catalyst is for example in the form of extrudates.

[0227] In one embodiment, the hydrocracking catalyst used in the second hydrocracking step comprises a hydro-dehydrogenating function including at least one Group VIII noble metal selected from palladium and platinum, alone or in a mixture. The Group VIII noble metal content is advantageously between 0.01 and 5 wt% and preferably between 0.05 and 3 wt%, the percentages being expressed as a wt% of oxides (PtO or PdO) relative to the total wt% of the catalyst.

[0228] According to another aspect of the invention, the hydrocracking catalyst further comprises one or more organic compounds containing oxygen and / or nitrogen and / or sulfur. Such a catalyst is often referred to as an "additized catalyst." Generally, the organic compound is chosen from a compound having one or more chemical functions selected from among a carboxyl group, alcohol, thiol, thioether, sulfone, sulfoxide, ether, aldehyde, ketone, ester, carbonate, amine, nitrile, imide, oxime, urea, and amide, or compounds including a furanic ring, or sugars.

[0229] The preparation of catalysts for the hydrogenation, hydrotreating, and hydrocracking steps is known and generally includes an impregnation step with Group VIII and Group VIB metals, when present, and possibly phosphorus and / or boron on the support, followed by drying and then possibly calcination. In the case of an additive catalyst, preparation is generally carried out by simple drying without calcination after the introduction of the organic compound. Calcination is understood here as heat treatment under a gas containing air or oxygen at a temperature of 200°C or higher. Before their use in a step of the process, catalysts are generally subjected to sulfidation to form the active species. The catalyst in step a) may also be a catalyst used in its reduced form, thus involving a reduction step in its pre- preparation.

[0230] The hydrogen-containing gas stream that feeds the hydrogenation, hydrotreating, and optionally hydrocracking reaction section can consist of hydrogen make-up and / or recycled hydrogen, particularly from separation step c) or fractionation step e) if this step is implemented. Preferably, an additional hydrogen-containing gas stream is advantageously introduced at the inlet of each reactor, particularly those operating in series, and / or at the inlet of each catalytic bed from the second catalytic bed of the reaction section. These additional gas streams are also called cooling streams. They allow for temperature control in the reactor, in which the reactions carried out are generally highly exothermic.

[0231] Said hydrotreated effluent or said hydrocarbon fraction(s) thus obtained by treatment according to the process of the invention of a pyrolysis oil, have a composition compatible with the specifications of an inlet feed for a steam cracking unit. In particular, the composition of the hydrotreated effluent or said hydrocarbon fraction(s) is preferably such that: - the total metallic content is less than or equal to 10.0 ppm by weight, preferably less than or equal to 2.0 ppm by weight, preferably less than or equal to 1.0 ppm by weight and preferably less than or equal to 0.8 ppm by weight, with: a silicon (Si) content of less than or equal to 5.0 ppm by weight, preferably less than or equal to 1 ppm by weight, and preferably less than or equal to 0.6 ppm by weight and an iron (Fe) content less than or equal to 200 ppb by weight, - the sulfur content is 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 and 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 olefinic compounds (mono- and di-olefins) is less than or equal to 5.0% by weight, preferably less than or equal to 2.0% by weight, 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.

[0232] The contents are given as relative weight concentrations, percentage (%) by weight, part(s) per million (ppm) by weight or part(s) per billion (ppb) by weight, relative to the total weight of the stream considered.

[0233] The process according to the invention therefore makes it possible to treat pyrolysis oils to obtain at least one effluent which can be injected, in whole or in part, into at least one steam cracking unit.

[0234] Heavy metal adsorption step (optional)

[0235] Any gaseous effluent and / or any liquid effluent from at least one of the separation steps b) and c) or the fractionation step e) may be subjected to an optional heavy metal adsorption step.

[0236] The gaseous effluents include, in particular, the first gaseous effluent from step b) and / or the second gaseous effluent from step c) and / or the third gaseous effluent from step e).

[0237] Liquid effluents include, in particular, the liquid effluent from step b) and / or the hydrocarbon effluent from step c) and / or at least one of the cuts from step e).

[0238] The optional adsorption step makes it possible to remove or reduce the amount of metallic impurities, in particular the amount of heavy metals such as arsenic, zinc, lead, and especially mercury, which may be present in said gaseous and liquid effluents. The metallic impurities may be present in the feedstock and / or form during the process steps. Their removal or reduction may be necessary, in particular, when at least part of said gaseous and liquid effluents is intended to be sent to a steam cracking step, either directly or after undergoing one or more optional additional steps such as the fractionation step (e). The specifications for metallic impurities at the inlet of the steam cracking step, particularly mercury, may necessitate such a step.

[0239] Thus, an optional step of adsorption of a gaseous effluent from steps b), c) and / or e) and / or of the liquid effluent from step b) and / or of the hydrocarbon effluent from step c) and / or at least one of the cuts from step e) is advantageously carried out in particular when at least one of these effluents or the feed includes respectively more than 20 ppb by weight, in particular more than 15 ppm by weight of metallic elements of heavy metals (As, Zn, Pb, Hg, ...), and in particular when at least one of these effluents or the feed includes respectively more than 10 ppm by weight of mercury, more particularly more than 15 ppm by weight of mercury.

[0240] Said optional adsorption step is advantageously carried out at a temperature between 20 and 150°C, preferably between 40 and 100°C, and at a pressure between 0.15 and 10.0 MPa abs, preferably between 0.2 and 1.0 MPa abs.

[0241] Said optional adsorption step can be implemented by any adsorbent known to those skilled in the art that allows the quantity of such contaminants to be reduced.

[0242] According to one variant, said optional adsorption step is implemented in an adsorption section operated in the presence of at least one adsorbent comprising a porous support and at least one active phase based on sulfur in elemental form or in the form of metallic sulfide.

[0243] The porous support may be chosen from among the families of aluminas, silica-aluminas, silicas, zeolites, and activated carbons. Advantageously, the porous support is alumina-based. The specific surface area of ​​the support is generally between 150 and 600 m² / g, preferably between 200 and 400 m² / g, and even more preferably between 150 and 320 m² / g. The specific surface area of ​​the adsorbent is measured by the BET method as described above.

[0244] The active phase is sulfur-based in elemental form or in the form of metallic sulfide, in particular a sulfide of a metal from the group selected from copper, molybdenum, tungsten, iron, nickel or cobalt.

[0245] Advantageously, the active phase of the adsorbent comprises between 1 and 70% by weight of sulfur relative to the total weight of the adsorbent, preferably between 2 and 25% and most preferably between 3 and 20%. The proportion by weight of metal relative to the total weight of the adsorbent is generally between 1 and 60%, preferably between 2 and 40%, most preferably between 5 and 30%, most preferably between 5 and 20%.

[0246] The residence time in the adsorption section is generally between 1 and 180 minutes.

[0247] Said adsorption section comprises at least one adsorption column, preferably comprising at least two adsorption columns, preferably between two and four adsorption columns, containing said adsorbent. When the adsorption section comprises two adsorption columns, one operating mode may be a so-called "swing" mode, in which one of the columns is in operation, while the other column is in reserve. Another operating mode is to have at least two columns operating in series in a switchable mode. (g) Steam cracking step (optional)

[0248] The hydrotreated effluent from hydrotreatment step d), or at least one of the liquid hydrocarbon cuts from the optional step e), may be sent in whole or in part to a steam cracking step g).

[0249] Advantageously, the gaseous effluent(s) from step c) of separation and / or e) of fractionation and containing ethane, propane and butane, may, in whole or in part, also be sent to step g) of steam cracking.

[0250] Said steam cracking step (g) is advantageously carried out in at least one pyrolysis furnace at a temperature between 700 and 900°C, preferably between 750 and 850°C, and at a relative pressure between 0.05 and 0.3 MPa. The residence time of The separation time for hydrocarbon compounds is generally less than or equal to 1.0 second (denoted s), preferably between 0.1 and 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, advantageously in the form of steam, is advantageously between 0.3 and 3.0 kg of water per kg of hydrocarbon compounds entering step (g). Preferably, the optional step (g) is carried out in several pyrolysis furnaces in parallel so as to adapt the operating conditions to the different streams feeding step (g), particularly those from step (e), and also to manage the tube decoking times. A furnace comprises one or more tubes arranged in parallel. A furnace can also refer to a group of furnaces operating in parallel. For example, a furnace may be dedicated to cracking the middle distillates cut.

[0251] The effluents from the various steam cracking furnaces are generally recombined before separation to form a final effluent. It is understood that step g) of steam cracking includes the steam cracking furnaces but also the substeps associated with steam cracking that are well known to those skilled in the art. These substeps may include, in particular, heat exchangers, columns and catalytic reactors, and recirculation to the furnaces. A column generally allows the effluent to be fractionated in order to recover at least a light fraction comprising hydrogen and compounds having 2 to 5 carbon atoms, and a fraction comprising pyrolysis gasoline, and possibly a heavier fraction. Columns allow the different constituents of the light fractionation fraction to be separated in order to recover at least an ethylene-rich cut (C2 cut) and a propylene-rich cut (C3 cut), and possibly a butene-rich cut (C4 cut).Catalytic reactors are particularly useful for hydrogenating C2, C3, and even C4 fractions, as well as pyrolysis gasoline. Saturated compounds, especially those with 2 to 4 carbon atoms, are advantageously recycled to steam cracking furnaces to increase overall olefin yields.

[0252] This step (g) of steam cracking makes it possible to obtain at least one effluent containing olefins comprising 2, 3 and / or 4 carbon atoms (i.e. C2, C3 and / or C4 olefins), at satisfactory levels, in particular greater than or equal to 30 wt%, especially greater than or equal to 40 wt%, or even greater than or equal to 50 wt% of total olefins comprising 2, 3 and 4 carbon atoms relative to the weight of the steam cracking effluent considered. These C2, C3 and C4 olefins can then be advantageously used as polyolefin monomers. Analytical methods used

[0253] The analytical methods and / or standards used to determine the characteristics of the different flows, in particular the load to be treated and the effluents, are known to The person skilled in the art. They are listed below for informational purposes. Other methods considered equivalent may also be used, including equivalent IP, EN, or ISO methods:

[0254] [Tables 1] Description Methods Density @15°C ASTM D4052 Sulfur Content ISO 20846 Nitrogen Content ASTM D4629 Acid Value ASTM D664 Bromine Value ASTM DI 159 Maleic Anhydride Value MAV Method (1) Oxygen Content Combustion + Infrared Paraffin Content UOP990-11 Naphthenes and Olefins Content UOP990-11 Aromatics Content UOP990-11 Halogen Content ASTM D7359 Chlorine Content ASTM D7536 Metal Content: ASTM D5185 P Fe Si Na B Simulated Distillation ASTM D2887 1. MAV method described in the article: C. Lôpez-Garcia 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. l,pp. 57-68. List of figures

[0255] Mention of the elements referenced in Figures 1 to 2 allows for a better understanding of the invention, without the latter being limited to embodiments by The specific features illustrated in Figures 1 to 2. The various embodiments presented can be used alone or in combination with each other, without any limitation on the combination. In the figures, the same reference numerals designate identical or analogous elements. [Fig 1]

[0256] Figure 1 shows a diagram of a particular embodiment of the process of the present invention, comprising:

[0257] - a step a) of hydrogenating a hydrocarbon feedstock obtained from the pyrolysis of plastics 1 preferably mixed with at least part of the recycled liquid effluent 7a from step b) and in the presence of a hydrogen-rich gas 2 and possibly an amine supplied by stream 3 and possibly a sulfur compound supplied by stream 4, carried out in at least one fixed-bed reactor comprising at least one hydrogenation catalyst, to obtain a hydrogenated effluent 5;

[0258] - a step b) of separating the hydrogenated effluent 5 carried out at high pressure and high temperature (HHPS) to obtain at least a first gaseous effluent 6, and a liquid effluent 7 of which a part 7a can be recycled upstream of step a),

[0259] - a separation step c) carried out at high pressure and low temperature (HPLS) and supplied by the first gaseous effluent 6 and the liquid effluent 7b from step b) and an aqueous solution 8 and enabling the production of at least a second gaseous effluent 9 comprising hydrogen, an aqueous effluent 10 containing dissolved salts, and a hydrocarbon effluent 11;

[0260] - a step d) of hydrotreating the hydrocarbon effluent 11 from step c), in presence of hydrogen 12 carried out in at least one fixed bed reactor comprising at least one hydrotreating catalyst, to obtain a hydrotreated effluent 13;

[0261] - optionally a step e) of fractionating the hydrotreated effluent 13 allowing to obtain at least a third gaseous effluent 14, a naphtha cut 15 (generally comprising compounds having a boiling point less than or equal to 175°C) and a middle distillates cut 16 (generally comprising compounds having a boiling point greater than 175°C).

[0262] Following step e), a portion of the naphtha 15 fraction can be sent to a steam cracking process (not shown). Another portion of the naphtha 15 fraction can feed into step a) hydrogenation and / or step d) hydrotreating (not shown). [Fig 2]

[0263] Figure 2 shows a diagram of another particular embodiment of the process of the present invention, which is based on the diagram in [Fig. 1]. This diagram includes, in particular, a hydrocracking step f) in which at least a portion of the cut middle distillates 16 from step e) feeds this hydrocracking step f) which is carried out in at least one fixed bed reactor comprising at least one hydrocracking catalyst and is supplied with hydrogen 17. The hydrocracking effluent 18 is recycled upstream of the separation step b).

[0264] Instead of injecting the amine 3 flux into the inlet of step a) hydrogenation, it is possible to inject it into the inlet of step d) hydrotreating, into the inlet of step c) separation, into the inlet of step f) hydrocracking when it is present, or not to inject it, depending on the characteristics of the feed.

[0265] Only the main stages, with the principal flows, are shown in Figures 1 and 2, to facilitate a better understanding of the invention. It is understood that all the equipment necessary for operation is present (tanks, pumps, heat exchangers, furnaces, columns, etc.), even if not shown. It is also understood that hydrogen-rich gas flows (make-up or recycle), as described above, can be injected at the inlet of each reactor or catalytic bed or between two reactors or two catalytic beds. Methods well known to those skilled in the art for hydrogen purification and recycling can also be implemented. Examples

[0266] Example 1 is a non-conforming example of the invention without separation of halogenated compounds between step a) of hydrogenation and step d) of hydrotreating.

[0267] Example 2 is an example according to the invention with separation of halogenated compounds by the hot separation steps b) and cold separation steps c) between the hydrogenation step a) and the hydrotreating step d).

[0268] The feed 1 treated in the process is a plastics pyrolysis oil (i.e. comprising 100% by weight of said plastics pyrolysis oil) having the characteristics indicated in Table 2.

[0269] Table 2: Load characteristics Description Methods Unit Charge Density @15°C ASTM D4052 g / cm3 0.835 Sulfur Content ISO 20846 ppm wt. 135 Nitrogen Content ASTM D4629 ppm wt. 2109 Acid Value ASTM D664 mg KOH / g 1.5 Bromine Value ASTM DI 159 g / 100g 80 Maleic Anhydride Value MAV Method mg / 100g 10 Olefin Content UOP990-11 % wt. 25 Chlorine Content ASTM D7536 ppm wt. 500 Silicon Content ASTM D5185 ppm wt. 92 Simulated Distillation ASTM D2887 0% °C 40 10% °C 98 30% °C 161 50% °C 232 70% °C 309 90% °C 394 100% °C 432

[0271] The charge 1 is subjected to a step a) of hydrogenation carried out in a fixed bed reactor and in the presence of hydrogen 2 and a NiMo on alumina type hydrogenation catalyst under the operating conditions indicated in Table 4 allowing to obtain a hydrogenated effluent 5.

[0272] At the end of step a) of hydrogenation, the conversion rates (= (initial concentration - final concentration) / initial concentration) observed are shown in Table 3.

[0273] Table 3: Species conversions during step a) hydrogenation

[0274] [Tables3] Chlorine conversion rate % >90 Diolefin conversion rate % >70 Olefin conversion rate % >70 Silicon retention % >85

[0275] According to Example 1 (non-conforming), the hydrogenated effluent 5 from step a) of hydrogenation is directly subjected, without separation, to a hydrotreating step d) carried out in a fixed bed in the presence of hydrogen 12 and a NiMo-on-alumina hydrotreating catalyst under the conditions shown in Table 4. The hydrotreated effluent is then introduced into a low-pressure, low-temperature separator vessel to remove dissolved gases. The properties of the hydrotreated effluent, free of dissolved gases, are shown in Table 4 and conform to the specifications of a feed for a steam cracker.

[0276] According to example 2 (conforming), the hydrogenated effluent 5 from step a) of hydrogenation is subjected to a step b) of separation (hot) at a pressure substantially identical to that of step a) and whose temperature is controlled at 300°C, allowing to obtain a gaseous effluent 6 and a liquid effluent 7.

[0277] The gaseous effluent 6 and the entire liquid effluent 7 are mixed and then subjected to a cold separation step c): a stream of water 8 is injected into the mixture, the final mixture reaches a temperature of 40°C in a cold HP vessel operating at a pressure substantially identical to that of step b), from which a hydrogen-rich gas fraction 9, an aqueous fraction 10, and the washed hydrocarbon effluent 11 are obtained. The hydrocarbon effluent 11 is then introduced by means of a pump to adjust the total pressure in the hydrotreating step d), which is carried out in a fixed bed and in the presence of hydrogen 12 to adjust the partial pressure of hydrogen, and a NiMo-on-alumina hydrotreating catalyst under the conditions shown in Table 4.

[0278] The hydrotreated effluent is then introduced into a low-pressure, low-temperature separator vessel to remove dissolved gases. The properties of the hydrotreated effluent, free of dissolved gases, are shown in Table 4 and conform to the specifications of a feedstock for a steam cracker. All or part of the hydrotreated effluent obtained can then be used in a steam cracking step to form olefins, which can be polymerized to produce recycled plastics.

[0279] Table 4: Operating conditions of steps a), b), c) and d) and properties of the hydrotreated effluent. Example 1 (non-compliant) Example 2 (compliant) Stage a) Hydrogenation Average temperature (WABT) °C 280 280 ppH2 at outlet of a) MPa 5.2 5.2 Total pressure at outlet of a) MPa 6.1 6.1 h2 / hc NmVm3 300 300 WH h1 0.64 0.64 Stage b) Hot separation T °C 300 Total pressure at outlet of b) MPa 5.8 Stage c) Cold separation T °C 40 Total pressure at outlet of c) MPa 5.2 Stage d) Hydrotreating Average temperature (WABT) °C 302 292 ppH2 at outlet of d) MPa 5.2 5.2 Total pressure at outlet of d) MPa 5.9 5.9 h2 / hc NmVm3 300 300 WH h1 0.76 0.76 Properties of the hydrotreated effluent N content ppm wds <5 <5 S content mg / kg < 10 < 10 Cl content mg / kg <5 <5 Metal content mg / kg < 1 < 1 Si content mg / kg < 1 < 1 Hg content ppb <5 <5

[0281] It is observed that to obtain a hydrotreated effluent meeting the specifications of a feed for a steam cracker, the average temperature of the hydrotreatment step d) according to example 2 is 10°C lower than that according to example 1.

[0282] When the catalyst deactivates during its catalytic cycle, the average temperature can be increased to compensate for the catalytic deactivation. The temperature increase is possible up to a temperature at which catalyst replacement becomes necessary (the end-of-cycle temperature).

[0283] When a temperature difference of 60°C is fixed between the beginning and end of the catalytic cycle, a deactivation rate of the catalyst and the duration of the catalytic cycle can be deduced.

[0284] Table 5 shows the effect on the cycle time of step d) of hydrotreatment with or without the separation steps b) and c). It can be seen that the cycle time according to example 1 is 12 months, while the cycle time according to example 2 is 30 months.

[0285] Table 5: Catalytic cycle time of the hydrotreating step with AT (SOR*-EOR**) = 60 °C Example 1 (non-compliant) Example 2 (compliant) Deactivation Rate °C / month 5 2 AT (SOR*-EOR**) °C 60 60 Tsor °C 302 292 TeOR °C 362 352 Cycle Time per month 12 30

[0287] (*) SOR (start of cycle or start of run according to Anglo-Saxon terminology)

[0288] (*) EOR (end of cycle or end of run according to Anglo-Saxon terminology)

Claims

1. Demands A process for treating a feedstock comprising a pyrolysis oil of plastics and / or tires comprising halogenated compounds, said process comprising: a) a hydrogenation step carried out in a hydrogenation reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrogenation catalyst, said hydrogenation reaction section being fed at least by said feedstock, optionally mixed with at least a portion of a liquid effluent from a separation step b), and a first gaseous stream comprising hydrogen, said hydrogenation reaction section being carried out at an average temperature between 140 and 400°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a hydrogenated effluent, b) a separation step, fed by the effluent from step a), said step being carried out at a temperature above the precipitation temperature of ammonium halides and at a pressure substantially identical to the pressure of step a) to obtain at least a first gaseous effluent and a liquid effluent, part of which may be recycled upstream of step a), (c) a separation step, fed by the first gaseous effluent and at least part of the liquid effluent from step (b), and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of ammonium halides and at a pressure substantially the same as or lower than the pressure of step (b), to obtain at least a second gaseous effluent, an aqueous effluent and a hydrocarbon effluent, (d) a hydrotreating step implemented in a hydrotreating reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrotreating catalyst, said hydrotreating reaction section being fed at least by said hydrocarbon effluent from step (c), optionally by at least a portion of the liquid effluent from step (b), and a second gaseous stream comprising hydrogen, said reaction section hydrotreating being carried out at an average temperature between 250 and 430°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a hydrotreated effluent, e) optionally a fractionation step of all or part of the hydrotreated effluent from step d), to obtain at least a third gaseous effluent, a naphtha cut and at least one middle distillates cut, f) optionally a hydrocracking step carried out in a hydrocracking reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrocracking catalyst,said hydrocracking reaction section being fed by at least a portion of said hydrotreated effluent from step d) and / or by at least a portion of the middle distillate cut from step e) and a third gas stream comprising hydrogen, said hydrocracking reaction section being operated at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly volumetric rate between 0.1 and 10.0 h⁻¹ to obtain a first hydrocracking effluent.

2. A method according to the preceding claim, wherein step b) is carried out at a temperature between 200 and 450°C and step c) is carried out at a temperature greater than or equal to 20°C and less than 200°C.

3. A process according to any one of the preceding claims, wherein at least part of the liquid effluent obtained in step b) is recycled upstream of step a).

4. A method according to any one of the preceding claims comprising the fractionation step e).

5. A method according to any one of the preceding claims comprising the hydrocracking step f).

6. A method according to any one of the preceding claims, wherein in step a) the hydrogen coverage is between 250 and 800 Nm3 of hydrogen per m3 of charge (Nm3 / m3).

7. A process according to any one of the preceding claims, comprising at least one pretreatment step aO) of the pyrolysis oil fraction of plastics and / or tires, said pretreatment step being carried out work upstream of step a) and includes an adsorption step and / or a filtration step and / or a centrifugation step and / or a decantation step and / or an electrostatic separation step and / or a washing step using an aqueous solution and / or a gas stripping step.

8. A process according to any one of the preceding claims, wherein the hydrotreated effluent from hydrotreatment step d), or at least one of the cuts from step e), is sent in whole or in part to a steam cracking step g) carried out in at least one pyrolysis furnace at a temperature between 700 and 900°C and at a pressure between 0.05 and 0.3 MPa relative.

9. A method according to any one of the preceding claims wherein the reaction section of step a) implements at least two reactors operating in switchable mode.

10. A method according to any one of the preceding claims wherein a stream containing an amine and / or a sulfur compound is injected upstream of step a).

11. A process according to any one of the preceding claims wherein said hydrogenation catalyst comprises a support selected from alumina, silica, silica-aluminas, magnesia, clays and mixtures thereof and a hydro-dehydrogenating function comprising either at least one element of Group VIII and at least one element of Group VIB, or at least one element of Group VIII.

12. A process according to any one of the preceding claims wherein said hydrotreating catalyst comprises a support selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures thereof, and a hydro-dehydrogenating function comprising at least one element from group VIII and / or at least one element from group VIB.

13. A process according to any one of the preceding claims, further comprising a second hydrocracking step f') carried out in a hydrocracking reaction section, employing at least one fixed-bed reactor having n catalytic beds, n being an integer greater than or equal to 1, each comprising at least one hydrocracking catalyst, said hydrocracking reaction section being fed with at least a portion of the first hydrocracking effluent from the first hydrocracking step f) and a gas stream comprising hydrogen, said hydrocracking reaction section being implemented at a temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly volumetric velocity between 0.1 and 10.0 h1, to obtain a second hydrocracked effluent.

14. A process according to any one of the preceding claims, wherein said hydrocracking catalyst comprises a support selected from halogenated aluminas, combinations of boron and aluminum oxides, amorphous silica-aluminas and zeolites and a hydro-dehydrogenating function comprising at least one metal from group VIB selected from chromium, molybdenum and tungsten, alone or in mixture, and / or at least one metal from group VIII selected from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum.

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