WO2023208636A1

WO2023208636A1 - Method for treating plastic pyrolysis oil including an h2s recycling step

Info

Publication number: WO2023208636A1
Application number: PCT/EP2023/059938
Authority: WO
Inventors: Jerome Bonnardot; Loïk GUENAULT; Floriane MALDONADO; Inigo Ribas Sanguesa; Martin SANTOS MARTINEZ; Sheyla CARRASCO HERNANDEZ; Wilfried Weiss; Marisa DE SOUSA DUARTE
Original assignee: IFP Energies Nouvelles; Repsol S.A.
Priority date: 2022-04-29
Filing date: 2023-04-17
Publication date: 2023-11-02
Also published as: FR3135090A1; TW202407084A

Abstract

The present invention relates to a method for treating a plastic pyrolysis oil, the method comprising: - hydrotreating the feedstock in the presence of hydrogen and a catalyst; - separating/washing the hydrotreated effluent in the presence of an aqueous solution to obtain at least a first aqueous effluent and a hydrotreated hydrocarbon effluent; - separating the H2S contained in the first aqueous effluent to obtain a gas phase containing the H2S and a second aqueous effluent, the gas phase containing the H2S can be at least partially recycled upstream of step b); - separating the NH3 contained in the second aqueous effluent to obtain a gas phase containing NH3 and a third aqueous effluent. The present invention makes it possible, by recycling the H2S coming from the method, to decrease the consumption of the sulphurising agent to maintain the catalysts in sulphide form in feedstocks having low sulphur content.

Description

METHOD FOR PROCESSING PLASTICS PYROLYSIS OIL INCLUDING AN H ₂ S RECYCLING STEP

TECHNICAL AREA

The present invention relates to a process for treating a plastic pyrolysis oil in order to obtain a hydrocarbon effluent which can be recovered in a gasoline, jet or diesel fuel storage unit or as feedstock for a steam cracking unit. More particularly, the present invention relates to a process for treating a feedstock resulting from the pyrolysis of plastic waste making it possible to recycle a gas phase containing H ₂ S resulting from the process in order to maintain the catalysts in sulphide form in the catalytic stages. of the process and therefore reduce the consumption of sulfurizing agent to be added.

PRIOR ART

Plastics from collection and sorting sectors can undergo a pyrolysis stage in order to obtain, among other things, pyrolysis oils. These plastic pyrolysis oils are generally burned to generate electricity and/or used as fuel in industrial or district heating boilers.

Another way of valorizing plastic pyrolysis oils is the use of these plastic pyrolysis oils as feedstock for a steam cracking unit in order to (re)create olefins, the latter being constituent monomers of certain polymers. . However, plastic waste is generally mixtures of several polymers, for example mixtures of polyethylene, polypropylene, polyethylene terephthalate, polyvinyl chloride, polystyrene. In addition, depending on the uses, plastics can contain, in addition to polymers, other compounds, such as plasticizers, pigments, dyes or even residues of polymerization catalysts. Plastic waste may also contain, in a minority, biomass from, for example, household waste. The treatment of waste on the one hand, in particular storage, mechanical treatments, sorting, pyrolysis, and also the storage and transport of pyrolysis oil on the other hand can also induce corrosion. As a result, the oils resulting from the pyrolysis of plastic waste include a lot of impurities, in particular diolefins, metals, in particular iron, silicon, or even halogenated compounds, in particular chlorine-based compounds, heteroelements. such as sulfur, oxygen and nitrogen, insolubles, at often high contents and incompatible with steam cracking units or units located downstream of steam cracking units, in particular the steam cracking processes. polymerization and selective hydrogenation processes. These impurities can generate operability problems and in particular problems of corrosion, coking or catalytic deactivation, or even problems of incompatibility in the uses of the target polymers. The presence of diolefins can also lead to problems of instability of the pyrolysis oil characterized by the formation of gums. Gums and insolubles possibly present in the pyrolysis oil can cause clogging problems in the processes.

In addition, during the steam cracking stage, the yields of light olefins sought for petrochemicals, in particular ethylene and propylene, strongly depend on the quality of the feeds sent to steam cracking. The BMCI (Bureau of Mines Correlation Index according to Anglo-Saxon terminology) is often used to characterize hydrocarbon cuts. This index, developed for hydrocarbon products derived from crude oils, is calculated from the measurement of the density and the average boiling temperature: it is equal to 0 for a linear paraffin and to 100 for benzene. Its value is therefore all the higher as the product analyzed has an aromatic condensed structure, naphthenes having a BMCI intermediate between paraffins and aromatics.

Overall, the yields of light olefins increase when the paraffin content increases and therefore when the BMCI decreases. Conversely, the yields of undesired heavy compounds and/or coke increase when the BMCI increases.

Document WO 2018/055555 proposes a global, very general and relatively complex recycling process for plastic waste, ranging from the very stage of pyrolysis of plastic waste to the steam cracking stage. The process of application WO 2018/055555 comprises, among other things, a step of hydrotreating the liquid phase resulting directly from pyrolysis, preferably under fairly advanced conditions, particularly in terms of temperature, for example at a temperature between 260 and 300°C, a step of separating the hydrotreatment effluent then a step of hydrodealkylation of the heavy effluent separated at a preferably high temperature, for example between 260 and 400°C.

Unpublished patent application FR 21/00.026 describes a process for treating a plastic pyrolysis oil aimed at reducing and/or eliminating the impurities contained in the pyrolysis oil in order to obtain an effluent compatible with a steam cracker. . The process comprises the following steps: a) the hydrogenation of said feed in the presence of at least hydrogen and at least one hydrogenation catalyst at an average temperature of between 140 and 340°C, the outlet temperature of the step a) is at least 15°C higher than the temperature in entry to step a), to obtain a hydrogenated effluent; b) the hydrotreatment of said hydrogenated effluent in the presence of at least hydrogen and at least one hydrotreatment catalyst, to obtain a hydrotreated effluent, the average temperature of step b) being greater than the average temperature of the step a); c) separation of the hydrotreated effluent in the presence of an aqueous flow, at a temperature between 50 and 370°C, to obtain at least one gaseous effluent, an aqueous liquid effluent and a hydrocarbon liquid effluent.

One way to eliminate the impurities contained in ex-plastic pyrolysis oils is to carry out hydrotreatment in the presence of catalysts which are active in sulphide form.

In the context of fillers containing plastic pyrolysis oil, we generally have fillers which are quite low in sulfur. However, a minimum PPH2S is necessary in the hydrotreatment reactor to maintain the catalysts in sulphide form and thus not reduce them. To maintain sufficient PPH2S in the reactor and taking into account that the feeds do not contain sufficient sulfur, a sulfurizing agent is generally, or even necessarily, added continuously, typically DMDS (dimethyl disulfide) to the feed. The sulfurizing agent decomposes very quickly into H2S by the action of the temperature and the hydrogen at the reactor inlet and therefore provides the quantity of H2S necessary to ensure a minimum and sufficient PPH2S.

After hydrotreatment, at least part of the H2S contained in the effluent forms ammonium sulfide salts ((NH^S) with the NHs generated by the hydrogenation of nitrogen compounds during hydrotreatment. Unlike to conventional fossil-type fillers, ex-plastic pyrolysis oils generally contain higher nitrogen contents than sulfur contents. These salts are generally eliminated by washing with water, followed by a (single) step stripping with water vapor of the aqueous effluent making it possible to obtain a purified aqueous effluent and a gas phase containing H2S and NH3 which are generally evacuated together at the top of the stripping column. The gas phase containing h^S and NH3 is then generally burned to form SO _X (sulfur oxides) and N ₂ or NO _X (nitrogen oxides).

The gas phase containing H2S and NH3 could be recovered and returned to the inlet of the hydrotreating unit in order to maintain the PPH2S in the reactor without adding a sulfurizing agent. But the NHs contained in this gas phase prevents this from being done because there would be a concentration of NHs in the recycling loop which would be detrimental to the operation of the unit. In addition, the presence of NHs lowers ppH2. The gas phase containing h S and NH3 cannot therefore be reused directly as a source of H2S to maintain the catalysts in sulphide form.

The present invention proposes a process for treating a feed comprising a plastic pyrolysis oil allowing the recycling of a phase containing only H2S to use it as a source of H ₂ S at the input of the catalytic units of the process by carrying out a separation, generally by stripping, in two stages making it possible to separate the H ₂ S from the NH ₃ . Indeed, the use of two stripping columns operating under different operating conditions makes it possible to separate the H ₂ S from the NH ₃ and thus recover:

- a gas phase containing H ₂ S which can be recycled at the inlet of the hydrotreatment unit;

- and a gas phase containing NH ₃ which can be burned or also recycled at the entrance to the hydrotreatment unit;

Two-step stripping for separating H2S from NH ₃ thus has the following advantages:

- The elimination of NH ₃ in the phase containing only h^S allows the recycling of H2S as an input to the catalytic units of the process;

- Strong minimization of sulfurizing agent consumption;

- Elimination of the gas phase containing NH ₃ by burning it is easier because it no longer contains H ₂ S forming SOx type pollutants;

- The gas phase containing NH ₃ can also be recycled as an input to the catalytic units, advantageously in stoichiometric quantities adapted to the formation of salts during step c) of separation/washing;

- Better compliance with environmental constraints in terms of SOx because the majority of the h^S is not burned (but on the contrary recycled in a loop);

- A reduction in hydrogen consumption in the hydrotreatment unit because the sulfurizing agent (DM DS) consumes hydrogen to decompose;

- Total elimination of NH ₃ in the gaseous effluent comprising hydrogen and/or light hydrocarbons from the head of the separation/washing section (step c) described below). In fact, the NH ₃ was captured in the form of ammonium sulphide in the aqueous effluent by the excess h^S which was recycled. The gaseous effluent thus freed from NH ₃ can thus be sent to a steam cracker so as to increase the overall olefin yield. SUMMARY OF THE INVENTION

More specifically, the invention relates to a process for treating a feedstock comprising a plastics pyrolysis oil, comprising: a) optionally a hydrogenation step implemented in a hydrogenation 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 hydrogenation catalyst, said hydrogenation reaction section being supplied at least by said feed and a gas flow comprising l 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 volume velocity between 0.1 and 10.0 h' ¹ , to obtain a hydrogenated effluent, b) a hydrotreatment step implemented in a hydrotreatment reaction section comprising at least one hydrotreatment catalyst, said hydrotreatment reaction section being supplied at least by the feed or said hydrogenated effluent from step a) and a gas flow comprising hydrogen, said hydrotreatment 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 volume velocity between 0.1 and 10.0 h' ¹ , to obtain a hydrotreated effluent, c) a separation step, supplied by the hydrotreated effluent from step b) and possibly by the hydrocracked effluent from step g) and an aqueous solution to obtain at least one gaseous effluent, a first aqueous effluent and a hydrocarbon effluent, d) a step of separating the H ₂ S contained in the first aqueous effluent to obtain a phase gas containing H ₂ S and a second aqueous effluent, said gas phase containing H ₂ S being optionally at least partly recycled upstream of step a) and/or step b) and/or step g), e) a step of separating the NH3 contained in the second aqueous effluent to obtain a gas phase containing NH3 and a third aqueous effluent, said gas phase containing NH3 optionally being at least partly recycled upstream of the step a) and/or step b) and/or step g), f) optionally a step of fractionating all or part of the hydrocarbon effluent from step c), to obtain at least one effluent gaseous and at least a first hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C and a second hydrocarbon cut comprising compounds having a boiling point greater than 175°C, g) optionally a hydrocracking step implemented in a hydrocracking 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 hydrocracking catalyst, said hydrocracking reaction section being fed by at least part of said hydrocarbon effluent from step c) and/or by at least part of the second hydrocarbon cut comprising compounds having a boiling point greater than 175°C from step f) and a gas flow 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 volume velocity between 0.1 and 10.0 ^h'1 to obtain a first hydrocracked effluent.

The present invention therefore relates to a process making it possible to purify an oil resulting from the pyrolysis of plastic waste of at least part of its impurities, which allows it to be hydrogenated 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 treatment in a steam cracking unit while being able to recycle the H2S resulting from the process continuously in order to minimize the consumption of sulfurizing agent. The injection of a sulfurizing agent remains necessary in particular at the start of the catalytic cycle, until the H ₂ S is formed to be separated in step d) and recycled upstream of step a) and/or step b) and/or step g), and/or upstream of step aO) of selective hydrogenation. Additional injections throughout the catalytic cycle may be necessary to compensate for natural loss. However, the fact of being able to recycle a gas phase containing H ₂ S without NH ₃ by the present invention makes it possible to considerably reduce the consumption of the sulfurizing agent.

Another advantage is the elimination of NH ₃ in the gaseous effluent comprising hydrogen and/or light hydrocarbons from the head of the separation/washing section (step c) by reaction with H ₂ S in excess recycled in the form of ammonium sulphide in the aqueous effluent. In other words, the NHs leaves in the form of salt in the aqueous effluent.

Another advantage of the invention is to prevent risks of blockage 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 significant quantities , diolefins, metals and halogenated compounds in plastic pyrolysis oil.

The process of the invention thus makes it possible to obtain a hydrocarbon effluent from a plastic pyrolysis oil freed at least in part from the impurities of the pyrolysis oil. of starting plastics, thus limiting operability problems, such as corrosion, coking or catalytic deactivation problems, which these impurities can cause, in particular in steam cracking units and/or in units located downstream of the processing units. steam cracking, in particular polymerization and hydrogenation units. The elimination of at least part of the impurities in the oils resulting from the pyrolysis of plastic waste will also make it possible to increase the range of applications of the target polymers, with incompatibilities of use being reduced.

According to a variant, said gas phase containing the H ₂ S resulting from step d) is at least partly recycled upstream of step a) and/or step b) and/or step g) .

According to one variant, the process comprises the hydrogenation step a).

According to a variant, the process comprises the fractionation step f).

According to one variant, the process comprises the hydrocracking step g).

According to a variant, step d) of separating the H ₂ S contained in the first aqueous effluent is carried out by stripping said effluent with a flow containing water vapor at a pressure of between 0.5 and 1 MPa and a temperature between 80 and 150°C.

According to a variant, step e) of separation of the NH ₃ contained in the second aqueous effluent is carried out by stripping said effluent with a flow containing water vapor at a pressure of between 0.1 and 0.5 MPa and a temperature between 80 and 150°C.

According to a variant, separation step c) comprises the following steps: c1) a separation step, fed by the hydrotreated effluent from step b), said step being carried out at a temperature between 200 and 450° C and at a pressure substantially identical to the pressure of step b) to obtain at least one gaseous effluent and one liquid effluent, part of which is optionally recycled upstream of step a) and/or step b) , c2) a separation step, supplied by the gaseous effluent from step c1) and another part of the liquid effluent from step c1) and an aqueous solution, said step being carried out at a temperature between between 20 and less than 200°C, and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent, a first aqueous effluent and a hydrocarbon effluent.

According to a variant, the process comprises at least one step aO) of pretreatment of the load comprising a plastic pyrolysis oil, optionally mixed with the hydrocarbon effluent from step c), said pretreatment step being implemented upstream of step a) and/or upstream of step b) and comprises a filtration step and/or a centrifugation step and/or an electrostatic separation step and/or a washing step using an aqueous solution and/or an adsorption step and/or a selective hydrogenation step.

According to a variant, the hydrocarbon effluent resulting from step c) of separation, or at least one of the two liquid hydrocarbon cuts resulting from step f), is sent in whole or in part to a steam cracking step h) carried out in at least one pyrolysis oven at a temperature between 700 and 900°C and at a pressure between 0.05 and 0.3 MPa relative.

According to a variant, said gas phase containing NH3 resulting from step e) is at least partly recycled upstream of step a) and/or step b) and/or step g).

According to a variant, a flow containing a nitrogen compound and/or a sulfur compound is injected upstream of step a) and/or upstream of step b).

According to a variant, said hydrogenation catalyst comprises a support chosen from alumina, silica, silica-aluminas, magnesia, clays and their mixtures 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.

According to a variant, said hydrotreatment catalyst comprises a support chosen from 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.

According to a variant, the process further comprises a second hydrocracking step g') implemented in a hydrocracking reaction section, using at least one fixed bed reactor having n catalytic beds, n being a greater integer or equal to 1, each comprising at least one hydrocracking catalyst, said hydrocracking reaction section being fed by at least part of the first hydrocracked effluent resulting from the first hydrocracking step g) 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 volume velocity between 0.1 and 10.0 h' ¹ , to obtain a second hydrocracked effluent.

According to a variant, said hydrocracking catalyst comprises a support chosen 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 VI B chosen from chromium, molybdenum and tungsten, alone or in a mixture, and/or at least one metal from group VIII chosen from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum.

The invention also relates to the product capable of being obtained, and preferably obtained by the process according to the invention.

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,

- including an iron element content less than or equal to 200 ppb by weight, and/or

- a silicon element content less than or equal to 5.0 ppm by weight, and/or

- a sulfur content less than or equal to 100 ppm by weight, and/or

- a nitrogen content less than or equal to 100 ppm by weight, and/or

- a chlorine element content less than or equal to 10 ppm by weight, and/or

- a mercury content less than or equal to 5 ppb by weight.

According to the present invention, the pressures are absolute pressures, also denoted abs., and are given in absolute MPa (or abs. MPa), unless otherwise indicated.

According to the present invention, the expressions "between ... and ..." and "between .... and ..." are equivalent and mean that the limit values of the interval are included in the range of values described . If this were not the case and the limit values were not included in the range described, such precision will be provided by the present invention.

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

In the following, particular and/or preferred embodiments of the invention can be described. They can be implemented separately or combined with each other, without limitation of combination when technically feasible.

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 (or VI 11 B) according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IIIPAC classification.

The metal content is measured by X-ray fluorescence. DETAILED DESCRIPTION

Load

According to the invention, a “plastic pyrolysis oil” is an oil, advantageously in liquid form at room temperature, resulting from the pyrolysis of plastics, preferably plastic waste originating in particular from collection and sorting sectors. It can also come from the pyrolysis of used tires.

It comprises in particular a mixture of hydrocarbon compounds, in particular 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 can comprise up to 70% by weight of paraffins, up to 90% by weight of olefins and up to 90% by weight of aromatics, it being understood that the sum of paraffins, olefins and aromatics is 100% weight of hydrocarbon compounds.

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/cm ³ , preferably between 0.75 and 0.95 g/cm ³ .

The plastic pyrolysis oil can comprise, and most often includes, additionally impurities such as metals, in particular iron, silicon, halogenated compounds, in particular chlorinated compounds. These impurities can be present in the plastic pyrolysis oil at high levels, for example up to 350 ppm by weight or even 700 ppm by weight or even 1000 ppm by weight of halogen elements (in particular chlorine) provided by halogenated compounds, and up to 100 ppm weight, or even 200 ppm weight of metallic or semi-metallic elements. Alkali metals, alkaline earths, transition metals, poor metals and metalloids can be assimilated to contaminants of a metallic nature, called metals or metallic or semi-metallic elements. In particular, the metals or metallic or semi-metallic elements, possibly contained in the oils resulting from the pyrolysis of plastic waste, include silicon, iron or these two elements. The plastic pyrolysis oil may also include other impurities such as heteroelements provided in particular by sulfur compounds, oxygenated compounds and/or nitrogen compounds, at contents generally lower than 27,000 ppm weight of heteroelements and preferably lower at 15500 ppm weight of heteroelements. The sulfur compounds are generally present in a content of less than 2000 ppm by weight and preferably less than 500 ppm by weight. The oxygenated compounds are generally present in a content of less than 15,000 ppm by weight and preferably less than 10,000 ppm by weight. Nitrogen compounds are generally present in a content of less than 10,000 ppm by weight and preferably less than 5000 ppm by weight. The plastic pyrolysis oil may also include other impurities such as heavy metals such as mercury, arsenic, zinc and lead, for example up to 100 ppb by weight or even 200 ppb by weight of mercury.

The feed for the process according to the invention comprises at least one plastic pyrolysis oil. Said load may consist solely of plastic pyrolysis oil(s). Preferably, said filler comprises at least 50% by weight, preferably between 70 and 100% by weight, of plastic pyrolysis oil relative to the total weight of the filler, that is to say preferably between 50 and 100% by weight. 100% by weight, preferably between 70% and 100% by weight of plastic pyrolysis oil.

The feed for the process according to the invention may comprise, in addition to the plastic pyrolysis oil or oils, a conventional petroleum feed or a feed resulting from the conversion of the biomass which is then co-treated with the pyrolysis oil. of plastics from the load.

The conventional petroleum feedstock can advantageously be a cut or a mixture of cuts of the naphtha, gas oil or vacuum gas oil type.

The load resulting from the conversion of the biomass can advantageously be chosen from vegetable oils, algal or algal oils, fish oils, used food oils, and fats of vegetable or animal origin; or mixtures of such fillers. Said vegetable oils can advantageously be crude or refined, totally or in part, and derived from plants chosen from rapeseed, sunflower, soya, palm, olive, coconut, copra, castor, cotton. , 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. Said animal fats are advantageously chosen from bacon and fats composed of residues from the food industry or from the catering industries. Frying oils, various animal oils such as fish oils, tallow, lard can also be used. The feedstock resulting from the conversion of the biomass can also advantageously be chosen from methyl esters of fatty acids of plant and/or animal origin or even methyl esters of fatty acids from used edible vegetable oils.

The feedstock resulting from the conversion of biomass can also be chosen from feedstocks originating from thermal or catalytic biomass conversion processes, such as oils which are produced from biomass, in particular lignocellulosic biomass, with various methods. liquefaction, such as liquefaction hydrothermal or pyrolysis. The term "biomass" refers to material derived from recently living organisms, which includes plants, animals and their by-products. The term “lignocellulosic biomass” refers to biomass derived from plants or their by-products. Lignocellulosic biomass is composed of carbohydrate polymers (cellulose, hemicellulose) and an aromatic polymer (lignin).

The load resulting from the conversion of the biomass can also advantageously be chosen from loads originating from the paper industry.

Plastic pyrolysis oil can come from thermal or catalytic pyrolysis treatment or be prepared by hydropyrolysis (pyrolysis in the presence of a catalyst and hydrogen).

Preprocessing (optional)

Said feed comprising a plastic pyrolysis oil, optionally mixed with the hydrocarbon effluent from step c) can advantageously be pretreated in at least one optional pretreatment step aO), prior to step a) of hydrogenation and /or hydrotreatment step b), to obtain a pretreated feedstock which feeds step a) and/or step b).

According to a variant, 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, possibly present in the load comprising a pyrolysis oil of plastics. This optional step aO) allows in particular the elimination of sediments which may form due to the unstable nature of pyrolysis oils and/or a compatibility problem between two different charges. Thus, an optional step aO) of pretreatment of the load comprising a plastic pyrolysis oil is advantageously carried out in particular when said load comprises 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 charge comprises 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. Likewise, an optional step aO) of pretreatment of the charge comprising a plastic pyrolysis oil is advantageously carried out in particular when said charge comprises more than 10 ppm weight, in particular more than 20 ppm weight, more particularly more than 50 ppm weight of chlorine.

Said optional pretreatment step aO) can be implemented by any method known to those skilled in the art making it possible to reduce the quantity of contaminants. It may in particular comprise a filtration step and/or a centrifugation step and/or an electrostatic separation step and/or a washing step using an aqueous solution and/or an adsorption step and /or a selective hydrogenation step.

When the optional pretreatment step aO) comprises a filtration step and/or a centrifugation step and/or an electrostatic separation step and/or a washing step using an aqueous solution and/or an adsorption step, it 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.

According to a variant, said optional pretreatment step aO) is implemented in an adsorption section operated in the presence of at least one adsorbent, preferably of the alumina type, having a specific surface area greater than or equal to 100 m ² /g , preferably greater than or equal to 200 m ² /g. The specific surface area of said at least adsorbent is advantageously less than or equal to 600 m ² /g, in particular less than or equal to 400 m ² /g. The specific surface area of the adsorbent is a surface area measured by the BET method, that is to say the specific surface area determined by nitrogen adsorption in accordance with the ASTM D 3663-78 standard established using the BRUNAUER-EMMETT method -TELLER described in the periodical 'The Journal of the American Chemical Society", 6Q, 309 (1938). Advantageously, said adsorbent comprises less than 1% by weight of metallic elements, preferably is 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 IUPAC classification). The residence time of the charge in the adsorption section is generally between 1 and 180 minutes.

Said adsorption section of optional step aO) comprises at least one adsorption column, preferably comprises at least two adsorption columns, preferably between two and four adsorption columns, containing said adsorbent. When the adsorption section comprises two adsorption columns, an operating mode can be an operation called "swing", according to the established Anglo-Saxon term, in which one of the columns is in line, i.e. i.e. in operation, while the other column is in reserve. When the absorbent of the online column is worn out, this column is isolated while the reserve column is put online, that is to say in operation. The spent absorbent can then be regenerated in situ and/or replaced by fresh absorbent so that the column containing it can be put back online once the other column has been isolated. Another mode of operation is to have at least two columns operating in series. When the absorbent of the column placed at the top is worn out, this first column is isolated and the used absorbent is either regenerated in situ or replaced by fresh absorbent. The column is then put back online in last position and so on. This operation is called permutable mode, or according to the English term “PRS” for Permutable Reactor System or even “lead and lag” according to the established Anglo-Saxon term. The combination of at least two adsorption columns makes it possible to overcome poisoning and/or possible and possibly rapid clogging of the adsorbent under the joint action of metal contaminants, diolefins, gums derived from diolefins and insolubles possibly present in the pyrolysis oil of plastics to be 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 making it possible to reduce the risks of clogging and therefore to avoid unit shutdown due to clogging, to control costs and to limit adsorbent consumption.

According to another variant, said optional pretreatment step aO) is implemented in a washing section with an aqueous solution, for example water or an acidic or basic solution. This washing section may include equipment making it possible to bring the load into contact with the aqueous solution and to separate the phases so as to obtain the pretreated load on the one hand and the aqueous solution comprising impurities on the other hand. Among this equipment, there may for example be a stirred reactor, a decanter, a mixer-decanter and/or a co- or counter-current washing column.

According to another variant, said optional pretreatment step aO) is implemented by filtration. The filtration step makes it possible to eliminate inorganic solids, sediments and/or fines contained in the load, in particular metals, metal oxides and metal chlorides. A filter is generally used whose size (for example the diameter or equivalent diameter) of the pores is less than 25 pm, preferably less than or equal to 10 pm, even more preferably less than or equal to 5 pm. According to another variant, a filter can be used whose pore size is less than 25 pm but greater than 5 pm. It is also possible to use a series of filters with different pore sizes, in particular a series of filters having decreasing pore sizes in the direction of flow of the charge. These filter media are well known for industrial uses. Cartridge filters and self-cleaning filters are suitable, for example. The dry extract can be measured for example by the Heptane Insolubles test, ASTM Method D-3279. The content of heptane insolubles should be reduced to less than 0.5% by weight, preferably less than 0.1%. According to a particular embodiment, the pretreatment step aO) by filtration comprises at least one filter whose pore size is less than 10 microns, and preferably greater than 5 pm, followed by a filtration system whose size of the pores is less than 2 pm and preferably less than 1 pm.

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

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

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

According to another variant, said optional pretreatment step aO) comprises a selective hydrogenation step. This selective hydrogenation step is advantageously implemented in a reaction section supplied at least by said feedstock, optionally pretreated by one or more pretreatments described above, and a gas flow comprising hydrogen, in the presence of at least a selective hydrogenation catalyst, at a temperature between 100 and 280°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volume velocity between 0.3 and 10.0 ^h'1 , to obtain a hydrogenated effluent.

This selective hydrogenation step is carried out under hydrogen pressure and temperature conditions making it possible to maintain said charge in the liquid phase and with a quantity of soluble hydrogen just necessary for selective hydrogenation of the diolefins present in the pyrolysis oil. It is advantageously carried out under milder conditions than hydrogenation step a). The selective hydrogenation of diolefins in the liquid phase thus makes it possible to avoid or at least limit the formation of "gums", that is to say the polymerization of diolefins and therefore the formation of oligomers and polymers, which can clog the or downstream reaction sections. Styrenic compounds, in particular styrene, possibly present in the filler can also behave like diolefins in terms of gum formation due to the fact that the double bond of the vinyl group is conjugated with the aromatic ring. Said selective hydrogenation step makes it possible to obtain a selectively hydrogenated effluent, that is that is to say an effluent with a reduced content of olefins, in particular diolefins and possibly styrene compounds.

Said reaction section implements selective hydrogenation, preferably in a fixed bed, in the presence of at least one selective hydrogenation catalyst, advantageously at an average temperature (or WABT as defined below for step a) d hydrogenation) between 100 and 280°C, preferably between 120 and 260°C, preferably between 130 and 250°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs, preferably between 2.0 and 8.0 MPa abs and at an hourly volume velocity (WH) between 0.3 and 10.0 ^h'1 , preferably between 0.5 and 5.0 ^h'1 .

The quantity of the gas flow comprising hydrogen (H2), supplying said reaction section of the selective hydrogenation step, is advantageously such that the hydrogen coverage is between 1 and 200 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ), preferably between 1 and 50 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ), preferably between 5 and 20 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ).

The hourly volume velocity (WH) and the hydrogen coverage are as defined below for hydrogenation step a).

The selective hydrogenation step is preferably carried out in a fixed bed. It can also be carried out in a bubbling bed or in a moving bed.

Advantageously, the reaction section of said selective hydrogenation step comprises between 1 and 5 reactors. According to a particular embodiment of the invention, the reaction section comprises between 2 and 5 reactors, which operate in permutable mode, called according to the English term "PRS" for Permutable Reactor System or even "lead and lag" as described in step a) of hydrogenation. According to a particularly preferred variant, the selective hydrogenation reaction section comprises two reactors operating in switchable mode.

Said selective hydrogenation catalyst is generally a catalyst as described in hydrogenation step a). It may or may not be identical to the catalyst of hydrogenation step a).

The content of impurities, in particular diolefins, of the hydrogenated effluent obtained at the end of the selective hydrogenation step is reduced compared to that of the same impurities, in particular diolefins, included in the process feed. The selective hydrogenation step generally makes it possible to convert at least 20% and preferably at least 30% of the diolefins contained in the initial charge. Advantageously, said optional pretreatment step aO) comprises a filtration step and/or a centrifugation step and/or an electrostatic separation step and/or a washing step using an aqueous solution and/or an adsorption step, followed by a selective hydrogenation step.

Said optional pretreatment step aO) can also optionally be supplied with at least part of the hydrocarbon effluent from step c) of the process and/or part of the first hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C from step f) and/or a part of the second hydrocarbon cut comprising compounds having a boiling point greater than 175°C from step f), as a mixture or separately of the load comprising a plastic pyrolysis oil. The recycling of at least part of the liquid effluent from step c) and/or at least part of one or more hydrocarbon effluents from step f) makes it possible in particular to increase sedimentation. and therefore, after possible filtration, to improve the pretreatment of the load.

Said optional pretreatment step aO) thus makes it possible to obtain a pretreated feed which then feeds step b) and/or step a) of hydrogenation when it is present.

Hydrogenation step a) (optional)

According to the invention, the process optionally comprises a hydrogenation step a) implemented in a hydrogenation reaction section, implementing at least one fixed bed reactor having n catalytic beds, n being an integer greater than or equal to to 1, each comprising at least one hydrogenation catalyst, said hydrogenation reaction section being supplied at least by said feed, optionally pretreated, and a gas flow comprising hydrogen, said hydrogenation reaction section being implemented 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 volume velocity between 0.1 and 10.0 h- ¹ , to obtain a hydrogenated effluent.

Step a) is in particular carried out under hydrogen pressure and temperature conditions making it possible to carry out the hydrogenation of the diolefins, possibly remaining after the optional step of selective hydrogenation and of the olefins at the start of the reaction section. hydrogenation while allowing, through a rising temperature profile, to carry out hydrodemetallation and hydrodechlorination, particularly at the end of the hydrogenation reaction section. A necessary quantity of hydrogen is injected so as to allow the hydrogenation of at least part of the diolefins and olefins present in the plastic pyrolysis oil, the hydrodemetallation of at least part of the metals, in particular the retention of silicon, and also the conversion of at least part of the chlorine (into HCl). The hydrogenation of diolefins and olefins thus makes it possible to avoid or at least limit the formation of "gums", that is to say the polymerization of diolefins and olefins and therefore the formation of oligomers and polymers, which can plug the reaction section of hydrotreatment step b). In parallel with hydrogenation, hydrodemetallation, and in particular the retention of silicon during step a), makes it possible to limit the catalytic deactivation of the reaction section of step b) of hydrotreatment. In addition, the conditions of step a) make it possible to convert at least part of the chlorine.

Temperature control is important in this step and must respond to an antagonistic constraint. On the one hand, the temperature at the inlet and throughout the hydrogenation reaction section must be sufficiently high to allow the hydrogenation of diolefins and olefins at the start of the hydrogenation reaction section. On the other hand, the inlet temperature of the hydrogenation reaction section must be sufficiently low to avoid deactivation of the catalyst. The hydrogenation reactions, in particular of a part of the olefins and diolefins, being highly exothermic, we then observe a rising temperature profile in the hydrogenation reaction section. This higher temperature at the end of said section makes it possible to carry out the hydrodemetallation and hydrodechlorination reactions. Thus, the temperature at the outlet of the reaction section of step a) is higher than the temperature at the inlet of the reaction section of step a), generally by at least 3°C, preferably by at least 5 °C.

The temperature in step a) whether it is the average temperature (WABT), the temperature at the entrance to the reaction section or even the rise in temperature in step a) between entry and exit of the reaction section can in particular be controlled by injection of a diluent in step a), preferably a recycle of part of the liquid effluent from step c) and/or at least part of one or more hydrocarbon effluents from step f), in particular by the recycling rate and/or by the temperature of the recycled effluent.

The temperature difference between the inlet and outlet of the reaction section of step a) is understood with injection of a gaseous (hydrogen) or liquid cooling flow, in particular a part of the liquid hydrocarbon effluent from step c).

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 therefore means excluding the use of a heating means ( oven, heat exchanger etc.). The inlet temperature of the reaction section of step a) is between 135 and 397°C, preferably between 240 and 347°C.

The temperature at the outlet of the reaction section of step a) is between 138 and 400°C, preferably between 243 and 350°C.

According to the invention, it is advantageous to carry out the hydrogenation of the diolefins and part of the hydrodemetallation reactions in the same step and at a temperature sufficient to limit the deactivation of the catalyst of step a) which is manifested by a reduction of the conversion of diolefins. This same step also makes it possible to benefit from the heat of hydrogenation reactions, in particular from a portion of the olefins and diolefins, so as to have a rising temperature profile in this step and thus being able to eliminate the need for a heating device. between the catalytic hydrogenation section and the catalytic hydrotreatment section.

Said reaction section carries out 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 an hourly volume velocity (WH) between 0.1 and 10.0 h' ¹ , preferably between 0.2 and 5.0 h' ¹ , and very preferably between 0.3 and 3.0 h ' ¹ .

According to the invention, the “average temperature” of a reaction section corresponds to the Weight Average Bed Temperature (WABT) according to the established Anglo-Saxon term, well known to those skilled in the art. The average temperature is advantageously determined as a function of the catalytic systems, equipment and configuration thereof used. The average temperature (or WABT) is calculated as follows:

WABT = (Tentree + T ₅₀ rtje)/2 with Tentree: the temperature of the effluent entering the reaction section and T _SO rtie: the temperature of the effluent leaving the reaction section. Unless otherwise stated, the “average temperature” of a reaction section is given at cycle start conditions.

The hourly volumetric velocity (WH) is defined here as the ratio between the hourly volumetric flow rate of the charge comprising the plastic pyrolysis oil, possibly pretreated, by the volume of catalyst(s). Hydrogen coverage is defined as the ratio of the volume flow rate of hydrogen taken under normal conditions of temperature and pressure in relation to the volume flow rate of “fresh” charge, that is to say the charge to be treated, possibly pretreated. , without taking into account a recycled fraction, and in particular without taking into account the liquid effluent from step c) recycled and/or step f), at 15°C (in normal m ³ , denoted Nm ³ , of H ₂ per m ³ of load).

The quantity of the gas flow comprising hydrogen (H ₂ ), supplying said reaction section of step a), is advantageously such that the hydrogen coverage is between 100 and 1500 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ), preferably between 200 and 1000 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ), preferably between 250 and 800 Nm ³ of hydrogen per m ³ of charge ( Nm ³ /m ³ ).

Hydrogen can come from a fossil source or a renewable source, for example from the gasification of plastic waste or produced by electrolysis.

Advantageously, the reaction section of said step a) comprises between 1 and 5 reactors, preferably between 2 and 5 reactors, and particularly preferably it comprises two reactors. The advantage of a hydrogenation reaction section comprising several reactors lies in optimized treatment of the feed, while making it possible to reduce the risks of clogging of the catalytic bed(s) and therefore to avoid stopping the unit due to to clogging.

According to a preferred variant, these reactors operate in permutable mode, called “PRS” for Permutable Reactor System or even “lead and lag”. The combination of at least two reactors in PRS mode makes it possible to isolate a reactor, unload the spent catalyst, recharge the reactor with fresh catalyst and put said reactor back into service without stopping the process. PRS technology is described, in particular, in patent FR2681871.

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

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

Advantageously, said hydrogenation catalyst comprises a support, preferably mineral, and a hydro-dehydrogenating function. According to a variant, the hydro-dehydrogenating function comprises in particular at least one element from group VIII, preferably chosen from nickel and cobalt, and at least one element from group VI B, preferably chosen from molybdenum and tungsten. According to this variant, the total content expressed in oxides of the metallic elements of groups VI B 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 MoO and WO3 respectively.

The weight ratio expressed as metal oxide between the metal (or metals) of group VI B relative to the metal (or metals) of group VIII is preferably between 1 and 20, and more preferably between 2 and 10.

According to this variant, 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 in nickel oxide NiO relative to the weight of said catalyst), and between 1% and 30% by weight of molybdenum, preferably between 3% and 20% by weight of molybdenum (expressed as molybdenum oxide MoOs relative to the weight of said catalyst ) on a preferably mineral support, preferably on an alumina support.

According to another variant, the hydro-dehydrogenating function comprises, and preferably consists of at least one element from group VIII, preferably nickel. According to this variant, the content of nickel oxides 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 preferably mineral support, preferably on an alumina support.

The support of said hydrogenation catalyst is preferably chosen from alumina, silica, silica-aluminas, magnesia, clays and their mixtures. Said support may contain doping compounds, in particular oxides chosen from boron oxide, in particular boron trioxide, zirconia, ceria, titanium oxide, phosphoric anhydride and a mixture of these oxides. Preferably, said hydrogenation catalyst comprises an alumina support, optionally doped with phosphorus and optionally boron. When the 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 the alumina and advantageously at least 0.001% relative to the total weight of the alumina. The alumina used can for example be a y (gamma) or q (eta) alumina.

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

Very preferably, step a) can use in addition to the hydrogenation catalyst(s) described above, in addition 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 relative to the 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 MoOs relative to the weight of said catalyst, on an alumina support. This catalyst with low metal content can preferably be placed upstream or downstream of the hydrogenation catalyst(s) described above, preferably upstream.

Preferably, step a) can implement upstream of the hydrogenation catalyst(s) at least one guard bed containing adsorbents of the alumina, silica-alumina, zeolite and/or activated carbon type possibly containing metals of the group VI B and/or VIII. It is also possible to use a series of guard beds with particles of different diameters, in particular a series of guard beds having decreasing diameters in the direction of the flow of the charge (also called “grading” according to Anglo-Saxon terminology). .

Said hydrogenation step a) makes it possible to obtain a hydrogenated effluent, that is to say an effluent with a reduced content of olefins, in particular diolefins, and metals, in particular silicon. The content of impurities, in particular diolefins, of the hydrogenated effluent obtained at the end of step a) is reduced compared to that of the same impurities, in particular diolefins, included in the process feed. Hydrogenation step a) generally makes it possible to convert at least 40%, and preferably at least 60% of the diolefins as well as at least 40%, and preferably at least 60% of the olefins contained in the initial charge. The heat released by the saturation of the double bonds makes it possible to raise the temperature of the reaction medium and to initiate the hydrotreatment reactions, in particular the elimination, at least in part, of other contaminants, such as for example silicon and chlorine. Preferably, at least 50%, and more preferably at least 75%, of the chlorine and silicon of the initial charge are eliminated during step a). The hydrogenated effluent obtained at the end of hydrogenation step a) is sent, preferably directly, to hydrotreatment step b). Hydrotreatment step b)

According to the invention, the treatment process comprises a hydrotreatment step b) implemented in a hydrotreatment reaction section comprising at least one hydrotreatment catalyst, said hydrotreatment reaction section being supplied at least by the feed, optionally pretreated in step aO), or said hydrogenated effluent from step a) and a gas stream comprising hydrogen, said hydrotreatment 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 volume velocity between 0.1 and 10.0 ^h'1 , to obtain a hydrotreated effluent.

Advantageously, step b) implements the hydrotreatment reactions well known to those skilled in the art, and more particularly hydrotreatment reactions such as the hydrogenation of aromatics, hydrodesulfurization and hydrodenitrogenation. Additionally, hydrogenation of the remaining olefins and halogenated compounds as well as hydrodemetallation are continued.

Said hydrotreatment reaction section is advantageously carried out at a pressure equivalent to that used in the reaction section of step a) of hydrogenation, and generally at an average temperature higher than that of the reaction section of step a) hydrogenation. Thus, said hydrotreatment reaction section is advantageously carried out at an average hydrotreatment 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 an hourly volume velocity (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 , so preferred between 0.2 and ^1h'1 . The hydrogen coverage in step b) is advantageously between 100 and 1500 Nm ³ of hydrogen per m ³ of fresh charge, and preferably between 200 and 1000 Nm ³ of hydrogen per m ³ of fresh charge, so preferred between 250 and 800 Nm ³ of hydrogen per m ³ of fresh charge. The definitions of mean temperature (WABT), WH and hydrogen coverage correspond to those described above.

The hydrotreatment step is preferably carried out in a fixed bed. It can also be carried out in a bubbling bed, in a driven bed or in a moving bed. When the hydrotreatment step is carried out in a bubbling bed, in an entrained bed or in a moving bed, an additional hydrotreatment step in a fixed bed can be carried out in the same ranges of operating conditions after that in an ebullient bed, in an entrained bed or in a moving bed, with or without intermediate separation of a gas flow. Preferably, the treatment process comprises a hydrotreatment step b) implemented in a hydrotreatment 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 hydrotreatment catalyst.

Said hydrotreatment reaction section is fed at least by the feed or said hydrogenated effluent from step a) and a gas flow comprising hydrogen, advantageously at the level of the first catalytic bed of the first reactor in operation. An injection of at least part of the feed or the hydrogenated effluent from step a) and/or at least part of hydrogen between the different catalytic beds is also possible. Optionally, the reaction section of said step b) can also be additionally supplied with at least part of the liquid effluent from step c) and/or with at least part of one of the effluents from step f).

Advantageously, said step b) is implemented in a hydrotreatment 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, catalyst(s) d hydrotreatment. 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.

When step b) is implemented in a hydrotreatment reaction section comprising several, preferably two reactors, these reactors can operate in series and/or in parallel and/or in permutable mode (or PRS) and/or in “swing” mode. The different possible operating modes, PRS mode (or lead and lag) and swing mode, are well known to those skilled in the art and are advantageously defined above.

In another embodiment of the invention, said hydrotreatment 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, so favorite between two and five.

In a particularly preferred mode, the hydrogenation reaction section of step a) comprises two reactors operating in switchable mode followed by the hydrotreatment reaction section of step b) which comprises a single fixed bed reactor. Advantageously, said hydrotreatment catalyst used in said step b) can be chosen from known catalysts for hydrodemetallation, hydrotreatment, silicon capture, used in particular for the treatment of petroleum cuts, and their combinations. Known hydrodemetallation catalysts are for example those described in patents EP 0113297, EP 0113284, US 5221656, US 5827421, US 7119045, US 5622616 and US 5089463. Known hydrotreatment catalysts are for example those described in patents EP 0113297, EP 0113284, US 6589908, US 4818743 or US 6332976. Known silicon capture catalysts are for example those described in patent applications CN 102051202 and US 2007/080099.

In particular, said hydrotreatment 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 chosen from the group consisting of nickel and cobalt, and/or at least one element from group VI B, preferably chosen from the group group consisting of molybdenum and tungsten. The total content expressed in oxides of the metallic elements of 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 MoOs and WO3 respectively. The weight ratio expressed as metal oxide between the metal (or metals) of group VIB relative to the metal (or metals) of group VIII is preferably between 1.0 and 20, preferably between 2.0 and 10 For example, the hydrotreatment reaction section of step b) of the process comprises a hydrotreatment 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 hydrotreatment catalyst, and between 1.0% and 30% by weight of molybdenum, preferably between 3.0% and 29% by weight of molybdenum, expressed as oxide of molybdenum MoOs relative to the total weight of the hydrotreatment catalyst, on a mineral support, preferably on an alumina support.

The support of said hydrotreatment catalyst is advantageously chosen from alumina, silica, silica-aluminas, magnesia, clays and their mixtures. Said support may also contain doping compounds, in particular oxides chosen from boron oxide, in particular boron trioxide, zirconia, ceria, titanium oxide, phosphoric anhydride and a mixture of these oxides. . Preferably, said hydrotreatment catalyst comprises an alumina support, preferably an alumina support doped with phosphorus and possibly boron. When the 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 the alumina and advantageously at least 0.001% relative to the total weight of the alumina. The alumina used can for example be a y (gamma) or (eta) alumina.

Said hydrotreatment catalyst is for example in the form of extrudates.

Advantageously, said hydrotreatment catalyst used in step b) of the process has a specific surface area greater than or equal to 250 m ² /g, preferably greater than or equal to 300 m ² /g. The specific surface area of said hydrotreatment catalyst is advantageously less than or equal to 800 m ² /g, preferably less than or equal to 600 m ² /g, in particular less than or equal to 400 m ² /g. The specific surface area of the hydrotreatment catalyst is measured by the BET method, that is to say the specific surface area determined by nitrogen adsorption in accordance with the ASTM D 3663-78 standard established from the BRUNAUER-EMMETT- method. TELLER described in the periodical 'The Journal of the American Chemical Society', 6Q, 309 (1938). Such a specific surface area makes it possible to further improve the elimination of contaminants, in particular metals such as silicon.

According to another aspect of the invention, the hydrotreatment 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 “additive catalyst”. Generally, the organic compound is chosen from a compound comprising one or more chemical functions chosen from a carboxylic function, alcohol, thiol, thioether, sulfone, sulfoxide, ether, aldehyde, ketone, ester, carbonate, amine, nitrile, imide, oxime, urea and amide or even compounds including a furan cycle or even sugars.

Preferably, step b) can implement upstream of the hydrogenation catalyst(s) at least one guard bed or a series of “grading” type guard beds as described above for step has).

Advantageously, hydrotreatment step b) allows the hydrogenation of at least 80%, and preferably all of the olefins remaining after hydrogenation step a), but also the conversion at least in part of other impurities present in the feed, such as aromatic compounds, metal compounds, sulfur compounds, compounds nitrogen compounds, halogenated compounds (notably chlorinated compounds), oxygenated compounds. Preferably, the nitrogen content at the outlet of step b) is less than 100 ppm by weight. Preferably, the sulfur content at the outlet of step b) is less than 100 ppm by weight. Step b) can also make it possible to further reduce the content of contaminants, such as that of metals, in particular the silicon content. Preferably, the metal content at the output of step b) 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.

Depending on the content of sulfur compounds in the initial charge to be treated, a flow containing a sulfurizing agent can be injected upstream of the optional pretreatment step aO), the optional hydrogenation step a) and/or the hydrotreatment step b) and/or upstream of one of the optional hydrocracking steps g) when present, preferably upstream of hydrogenation step a) and/or step b) hydrotreatment in order to ensure a sufficient quantity of sulfur to form the active species of the catalyst (in sulphide form).

This activation or sulfidation step is carried out by methods well known to those skilled in the art, and advantageously under a sulfo-reducing atmosphere in the presence of hydrogen and hydrogen sulfide. The sulfurizing agents are preferably H2S gas, elemental sulfur, CS2, mercaptans, sulfides and/or polysulfides, hydrocarbon cuts with a boiling point below 400°C containing sulfur compounds or any other compound containing sulfur. used for the activation of hydrocarbon feeds with a view to sulfiding the catalyst. Said sulfur-containing compounds are advantageously chosen from alkyl disulfides such as for example dimethyl disulfide (DMDS), alkyl sulfides, such as for example dimethyl sulfide, thiols such as for example n- butyl mercaptan (or 1-butanethiol) and polysulphide compounds of the tertiononyl polysulphide type. The catalyst can also be sulfurized by the sulfur contained in the feed to be desulfurized. Preferably, the catalyst is sulfurized in situ in the presence of a sulfurizing agent and a hydrocarbon filler. Very preferably, the catalyst is sulphurized in situ in the presence of the additive charge of dimethyl disulphide.

The injection of a sulfurizing agent is particularly necessary at the start of the catalytic cycle, the time for the h^S to be formed to be separated in step d) and recycled upstream of step a) and/or step b) and/or step g), or even upstream of the selective hydrogenation step of the pretreatment aO). Additional injections throughout the catalytic cycle may be necessary to compensate for natural loss. However, the fact being able to recycle a gas phase containing H2S without NH3 by the present invention makes it possible to considerably reduce the consumption of the sulfurizing agent.

Step c) separation

According to the invention, the treatment process comprises a separation step c), advantageously implemented in at least one washing/separation section, supplied at least by the hydrotreated effluent from step b), and optionally by the hydrocracked effluent from optional steps g) and g'), and an aqueous solution, to obtain at least one gaseous effluent, a first aqueous effluent and a hydrocarbon effluent.

The gaseous effluent obtained at the end of step c) advantageously comprises hydrogen, preferably comprises at least 80% volume, preferably at least 85% volume, of hydrogen. Advantageously, said gaseous effluent can at least partly be recycled to the stages a) hydrogenation and/or b) hydrotreatment and/or g) hydrocracking, the recycling system possibly comprising a purification section.

The gaseous effluent can also be subject to additional separation(s) with a view to recovering at least one gas rich in hydrogen and/or light hydrocarbons, in particular ethane, propane and butane, which can advantageously be sent separately or mixed into one or more ovens of steam cracking step h) so as to increase the overall yield of olefins.

The hydrocarbon effluent from separation step c) is sent, in part or in full, either directly to the inlet of a steam cracking unit, or to an optional fractionation step f). Preferably, the hydrocarbon liquid effluent is sent, in part or in full, to a fractionation step f).

The first aqueous effluent obtained at the end of step c) advantageously comprises ammonium salts and/or hydrochloric acid as well as dissolved H ₂ S and NH ₃ .

This separation step c) makes it possible in particular to eliminate the ammonium chloride salts, which are formed by reaction between the chloride ions, released by the hydrogenation of the chlorinated compounds in HCl form, particularly during steps a) and b). then dissolution in water, and the ammonium ions, generated by the hydrogenation of the nitrogen compounds in the form of NH3 in particular during step b) and/or provided by injection of an amine then dissolution in water, and thus to limit the risks of blockage, in particular in the transfer lines and/or in the sections of the process of the invention and/or the transfer lines towards the steam cracker, due to the precipitation of ammonium chloride salts. He also allows the elimination of hydrochloric acid formed by the reaction of hydrogen ions and chloride ions. This step c) also makes it possible to eliminate the ammonium sulfide salts ((NH4)2S) which form by reaction between the H2S resulting from the hydrodesulfurization of the sulfur compounds and the NH3.

Depending on the content of chlorinated compounds in the initial charge to be treated, a flow containing a nitrogen compound such as an amine, for example monoethanolamine, diethanolamine and/or monodiethanolamine can be injected upstream of the step of selective hydrogenation of pretreatment aO) and/or upstream of hydrogenation step a) and/or between hydrogenation step a) and hydrotreatment step b) and/or between step g) hydrocracking and step c) of separation, preferably upstream of step a) of hydrogenation in order to ensure a sufficient quantity of ammonium ions to combine the chloride ions formed during the hydrotreatment step, thus making it possible to limit the formation of hydrochloric acid and thus to limit corrosion downstream of the separation section.

According to a variant, said gas phase containing the NH3 resulting from step e) can also be used as a nitrogen compound.

Advantageously, separation step c) comprises an injection of an aqueous solution, preferably an injection of water, into the hydrotreated effluent resulting from step b), or the hydrocracked effluent resulting from step g ) optional, upstream of the washing/separation section, so as to dissolve at least partly ammonium chloride salts and/or hydrochloric acid and thus improve the elimination of chlorinated impurities and reduce the risks of blockages due to accumulation of ammonium chloride salts.

Separation step c) is advantageously carried out at a temperature between 20 and 450°C, preferably between 100 and 440°C, preferably between 200 and 420°C. It is important to operate in this temperature range (and therefore not to cool the hydrotreated effluent too much) as there is a risk of blockage in the lines due to the precipitation of ammonium chloride salts. Advantageously, separation step c) is carried out at a pressure close to that implemented in steps a) and/or b), preferably between 1.0 and 10.0 MPa, so as to facilitate the recycling of 'hydrogen.

The separation step can advantageously be implemented 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 columns stripping gas, for example a gas flow rich in hydrogen. The washing/separation section of step c) can at least least partly be carried out in common or separate washing and separation equipment.

In a possible embodiment of the invention, separation step c) comprises the injection of an aqueous solution into the hydrotreated effluent resulting from step b), followed by the washing/separation section advantageously comprising a separation phase making it possible to obtain at least a first aqueous effluent loaded with ammonium salts, a washed liquid hydrocarbon effluent and a partially washed gaseous effluent. Said first aqueous effluent loaded with ammonium salts and the washed liquid hydrocarbon effluent can then be separated in a settling flask in order to obtain said hydrocarbon effluent and said first aqueous effluent. Said partially washed gaseous effluent can in parallel be introduced into a washing column where it circulates against the current of an aqueous flow, preferably of the same nature as the aqueous solution injected into the hydrotreated effluent, which makes it possible to eliminate at least in part, preferably in full, of the hydrochloric acid contained in the partially washed gaseous effluent and thus obtain said gaseous effluent, preferably comprising essentially hydrogen, and an acidic aqueous stream. Said first aqueous effluent from the settling flask can optionally be mixed with said acidic aqueous stream, and be used, possibly mixed with said acidic aqueous stream in a water recycling circuit to supply said solution to step c) of separation. aqueous flow upstream of the washing/separation section and/or in said aqueous flow in the washing column. Said water recycling circuit may include a top-up of water and/or a basic solution and/or a purge allowing the dissolved salts to be evacuated.

In another possible embodiment of the invention, separation step c) may advantageously comprise a “high pressure” washing/separation section which operates at a pressure close to the pressure of step a) of hydrogenation and/or step b) of hydrotreatment and/or step g) of optional hydrocracking, preferably between 1.0 and 10.0 MPa, in order to facilitate the recycling of hydrogen. This possible "high pressure" section of step c) can be supplemented by a "low pressure" section, in order to obtain a hydrocarbon effluent devoid of part of the gases dissolved at high pressure and intended to be treated directly in a steam cracking process or optionally be sent to fractionation step f).

Variant step c) of separation in two stages

In a possible embodiment of the invention, separation step c) comprises the following steps: c1) a separation step, fed by the hydrotreated effluent from step b), said step being carried out at a temperature between 200 and 450°C and at a pressure substantially identical to the pressure of step b) to obtain at least one gaseous effluent and one liquid effluent, part of which can be recycled upstream of step a) and/or step b), c2) a step separation, supplied by the gaseous effluent from step c1) and another part of the liquid effluent from step c1) and an aqueous solution, said step being carried out at a temperature between 20 and below 200°C, and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent, a first aqueous effluent and a hydrocarbon effluent.

Step c1)

According to the invention, the treatment process may comprise a separation step c1), supplied by the hydrotreated effluent from step b), said step being carried out at a temperature between 200 and 450°C and at a pressure substantially identical to the pressure of step b) to obtain at least one gaseous effluent and one liquid effluent, part of which can be recycled upstream of step a) and/or step b).

The separation step c1) is a so-called high pressure or medium pressure separation step at high temperature, also known to those skilled in the art under the name HHPS (for “Hot High Pressure Separator” according to Anglo-Saxon terminology). Thus, this step c1) preferably uses a so-called “high pressure hot” separator, the pressure being substantially equal to the operating pressure of step b). By “pressure substantially equal to the pressure of step b)” is meant the pressure of step b) with a pressure difference of between 0 and 1 MPa, preferably of between 0.005 and 0.3 MPa, and so particularly preferred between 0.01 and 0.2 MPa relative to the pressure of step b). Preferably, the pressure of step c1) is the pressure of step b) reduced by pressure losses.

The temperature at which the separation is carried out is between 200 and 450°C, preferably between 220 and 330°C, and particularly preferably between 240 and 300°C. According to a preferred variant, and in view of recovering the most calories, the separation is carried out at the highest possible temperature but less than or equal to the outlet temperature of step b) which makes it possible to avoid or limit reheating (and therefore a need for calories) of the effluent from step b). According to another variant, the effluent from step b) can be heated or cooled before separation.

This separation step c1) can advantageously be implemented 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 or these separator(s) (balloon(s)) and/or columns can optionally be powered by a stripping gas, for example a hydrogen-rich gas stream. Preferably, step c) is implemented with a single separator (balloon).

Part of the liquid effluent can be recycled upstream of step a) and/or step b) and/or upstream of the selective hydrogenation step of pretreatment aO). Recycling part of the product obtained towards or upstream of at least one of the reaction stages advantageously makes it possible on the one hand to dilute the impurities and on the other hand to control the temperature in the reaction stage(s) ( s), in which the reactions involved can be strongly exothermic.

Advantageously, the quantity of liquid effluent from step c) recycled, that is to say the fraction of product obtained recycled, is adjusted so that the weight ratio between the recycle stream from step c) and the feed comprising a plastic pyrolysis oil, that is to say 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 liquid effluent from step c) recycled is adjusted so that the weight ratio between the recycle stream and the load comprising a plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably between 0.2 and 5. This recycling rate makes it possible in particular to control the rise in temperature in step a). Indeed, when the recycling rate is high, the dilution rate of the feed is high, and the temperature rise at the start of the reaction section of step a), in particular due to the hydrogenation reactions of the diolefins, is thus controllable by the dilution effect.

Separation at high pressure and high temperature makes it possible, on the one hand, to maximize energy recovery by hot recycling of part of the liquid effluent. Indeed, the energy to reach the inlet temperature necessary in step a) and/or step b) can at least partly be provided by the heat of part of the liquid effluent from the step c) and also makes it possible to reduce or even eliminate possible preheating by direct heating of the load above a temperature above 200°C to avoid the formation of gums. In addition, recycling at least part of the liquid effluent at high pressure saves energy for its pressurization in step a) and/or step b).

The separation at high pressure and high temperature also makes it possible to minimize the quantity of light fraction (hydrocarbon cut comprising compounds having a point boiling point less than or equal to 175°C or naphtha) contained in the liquid effluent recycled in step a) and/or step b). At this temperature, almost all of the light fraction of the effluent (naphtha) leaves as gaseous effluent towards the separation/washing step c2) while in the liquid phase we have mainly the heavy fraction of the charge (cutting). hydrocarbon containing compounds having a boiling point greater than 175°C or middle distillates). In this way, the ppH ₂ is favored in step a) and/or in step b) because the light fraction (naphtha) could partially vaporize and lower the ppH ₂ if it was not at least in part eliminated during separation at high pressure and high temperature. The elimination of the light fraction comprising the naphtha can optionally be increased by a slight expansion upstream of at least one separator implemented in step c1) even if this implementation is not preferred due to the energy loss linked to relaxation. Another option for increasing the elimination of the light fraction comprising naphtha may consist of carrying out stripping, for example by injecting a gas rich in hydrogen in step c1).

According to a preferred variant, at least part of the hydrotreated liquid effluent from step c) can advantageously be either cooled, or pre-heated, if necessary, or kept at the same temperature as at the outlet of step c) separation, before being advantageously recycled upstream of step a) of hydrogenation and/or step b) of hydrotreatment, depending on the temperature and the flow rate of feed and hydrogen, so as to so that the temperature of the incoming flow, comprising said feed mixed with at least part of said liquid effluent from step c) and a hydrogen-rich gas, is between 140 and 430°C, preferably between 220 and 350°C, and particularly preferably between 260 and 330°C.

In the case where at least part of the liquid effluent from separation step c) is pre-heated before being recycled upstream of step a) and/or step b), said effluent optionally passes through at least one exchanger and/or at least one oven before being recycled upstream of step a) and/or step b), so as to adjust the temperature of said recycled liquid effluent.

In the case where at least part of the liquid effluent from separation step c) is cooled before being recycled upstream of step a) and/or step b), said effluent possibly passes in at least one exchanger and/or at least one air cooler before being recycled upstream of step a) and/or step b), so as to adjust the temperature of said recycled liquid effluent.

The implementation of the recycling upstream of step a) of hydrogenation and/or upstream of step b) of hydrotreatment, of at least part of the liquid effluent resulting from step c) Who can be either cooled or preheated, if necessary, or kept at the same temperature as at the outlet of separation step c), therefore makes it possible to adjust the temperature of the incoming flow in step a) and/or at l step b), as needed.

According to a variant, the feed, before being mixed with at least part of the effluent from step c), can be preheated by direct heating to a temperature of up to 200°C, preferably up to 200°C. at 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 clogging and an increase in the pressure loss of the load heating system as well. than the catalyst bed(s). Heating the load to a temperature above 150°C, preferably above 180°C, and particularly preferably above 200°C is preferably carried out by indirect heating by at least one part of the effluent from step c). Thus, the temperature rise above 150°C, preferably above 180°C and particularly preferably above 200°C of the charge is caused by mixing with a hotter liquid, and not by contact with a heated wall. According to another variant, the heating of the charge to a temperature above 150°C, preferably above 180°C, and particularly preferably above 200°C is carried out by ovens or exchangers sized to have a very low wall temperature compared to the temperature of the load, for example an electric oven.

According to another variant, the load is entirely heated by indirect heating by at least part of the effluent from step c). In this case, the load is not preheated before being mixed with at least part of the effluent from step c).

According to another variant, the load is not heated by indirect heating by at least part of the effluent from step c). In this case, the load and the part of the effluent from step c) recycled are mixed, the part of the effluent from step c) recycled having substantially the same temperature or a lower temperature than the load .

Another heating flow advantageously consists of a gaseous effluent rich in hydrogen coming from the hydrogen makeup and/or the gaseous effluent resulting from separation step c). At least part of this hydrogen-rich gaseous effluent coming from the hydrogen makeup and/or the gaseous effluent from separation step c) is advantageously injected mixed with at least part of the liquid effluent from step c) or separately, upstream of step a) and/or step b). The gaseous flow rich in hydrogen can therefore be advantageously either pre-heated in mixture with at least part of the liquid effluent or pre-heated separately before mixing preferably by possible passage through at least one exchanger and/or at least one oven or any other heating means known to those skilled in the art.

Step c2)

According to the invention, the treatment process comprises a separation step c2), advantageously implemented in at least one washing/separation section, supplied by the first gaseous effluent and another part of the liquid effluent from the step c1) and an aqueous solution, said step being carried out at a temperature between 20 and less than 200°C, and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent , a first aqueous effluent and a hydrocarbon effluent.

The separation step c2) is implemented in at least one so-called high pressure or medium pressure separator tank at low temperature, also known to those skilled in the art under the name CHPS (for “Cold High Pressure Separator” according to the terminology Anglo-Saxon). Thus, this step c2) preferably uses a so-called “high pressure cold” separator, the pressure being substantially equal to the operating pressure of step b). By “pressure substantially equal to the pressure of step b)” is meant the pressure of step b) with a pressure difference of between 0 and 1 MPa, preferably of between 0.005 and 0.3 MPa, and so particularly preferred between 0.01 and 0.2 MPa relative to the pressure of step b). Preferably, the pressure of step c2) is the pressure of step b) reduced by pressure losses. The fact of carrying out at least part of the separation step c2) at a pressure substantially identical to the operating pressure of step b) also facilitates the recycling of hydrogen.

Separation step c2) can also be carried out at a pressure lower than the pressure of step b).

The separation step c2) may also comprise 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 an identical or lower temperature. and at a lower pressure at each separation step from the previous step c2).

The temperature at which the separation of step c2) is carried out is between 20 and less than 200°C, preferably between 25 and 120°C, and particularly preferably between 30 and 70°C. It is important to operate in this temperature range (and therefore not to cool the hydrotreated effluent too much) as there is a risk of blockage in the lines due to the precipitation of ammonium chloride salts.

The washing/separation section of step c2) can at least partly be carried out in common or separate washing and separation equipment, this equipment being well known (separator flasks which can be operated at different pressures and temperatures, pumps, exchangers heat pumps, washing columns, etc.). Separation step c2) can in particular be carried out as step c) described above.

When one (or two) hydrocracking stages are present (described below), this stage c2) can additionally be supplied by at least part of the hydrocracked effluent resulting from an optional hydrocracking stage g) .

At least part of the hydrocarbon effluent from step c2) can be recycled as liquid quench (or quench according to Anglo-Saxon terminology) upstream of step a and/or step b) and/or step g). The injection of the hydrocarbon effluent from step c2) can be carried out at the first catalytic bed of the reaction section of step a) and/or step b) and/or step g) or between the different catalytic beds of each section. When the hydrogenation reaction section of step a) comprises two reactors operating in switchable mode, at least part of the hydrocarbon effluent from step c2) can be recycled between the two reactors.

Step d) of separation of the FkS contained in the first aqueous effluent

According to the invention, the treatment process comprises a step d) of separating the H2S contained in the first aqueous effluent to obtain a gas phase containing the H2S and a second aqueous effluent, said gas phase containing the H ₂ S is preferably at least partly recycled upstream of step a) and/or step b) and/or step g).

Step d) of separating the H ₂ S contained in the first aqueous effluent is advantageously carried out by stripping using a gas flow, preferably inert, in a stripping column. A stripping column is a distillation column in which a gas flow, preferably inert, is injected at the bottom of the column. The gas phase containing h^S is recovered at the top of the column and a second aqueous effluent is recovered at the bottom of the column. The inert gas stream may be hydrogen, nitrogen or steam. Preferably, step d) is carried out by steam stripping.

Stripping using a gas flow, preferably inert, makes it possible to obtain a very low content of H2S dissolved in the second aqueous effluent at the bottom of the stripping column. Step d) of separating the H2S is generally carried out at a pressure between 0.5 and 1.5 MPa, preferably between 0.5 and 1 MPa, and particularly preferably between 0.6 and 1.5 MPa. 0.9 MPa.

Stripping is generally carried out at a temperature between 80 and 150°C, preferably between 120 and 145°C (at the top and bottom of the column respectively).

The flow rate of the inert gas flow is generally such that the ratio between the flow rate of the inert gas flow expressed in normal m ³ per hour (Nm ³ /h) and the flow rate of the first aqueous effluent to be treated expressed in m ³ per hour at standard conditions (15°C, 0.1 MPa) is between 50 and 600 Nm ³ /m ³ , preferably between 200 and 400 Nm ³ /m ³ . By normal m ³ we mean the quantity of gas in a volume of 1 m ³ at 0°C and 0.1 MPa.

The stripping carried out under these operating conditions makes it possible in particular to separate H ₂ S from said first aqueous effluent without entraining the NH ₃ (which remains mainly in the aqueous phase).

According to a variant, in step d) a part of the gas phase at the top of the stripping column comprising the H ₂ S and the inert gas flow (steam) is condensed and is preferably reinjected at least in part as as liquid reflux in the upper part of the stripping column. Condensation is generally carried out by cooling to a temperature between 30 and 65°C, for example by cold water.

Liquid reflux makes it possible to control/reduce the temperature at the top of the stripping column.

Said gas phase containing the H ₂ S withdrawn at the top of the stripping column is at least partly recycled upstream of step a) and/or step b) and/or even upstream of step g) hydrocracking and/or upstream of the selective hydrogenation step of the pretreatment aO) when they are present, in order to act as a sulfurizing agent for the catalyst(s). Before recycling it can undergo at least one additional purification step, for example contact with liquid or washing with amines.

According to another variant, step d) of separating the H ₂ S contained in the first aqueous effluent can also be carried out by liquid/liquid extraction in which an inert or reactive solvent is brought into contact with the aqueous effluent.

Step e) of separation of the NH ₃ contained in the second aqueous effluent

According to the invention, the treatment process comprises a step e) of separating the NH ₃ contained in the second aqueous effluent from step d) to obtain a gas phase containing NH ₃ and a third aqueous effluent, said gas phase containing the NH3 preferably being at least partly recycled upstream of step a) and/or step b) and/or step g).

Step e) of separating the NH3 contained in the second aqueous effluent is advantageously carried out by stripping using an inert gas flow in a stripping column. The gas phase containing NH ₃ is recovered at the top of the column and a third aqueous effluent is recovered at the bottom of the column. The inert gas stream may be hydrogen, nitrogen or steam. Preferably, step e) is carried out by steam stripping.

Stripping using an inert gas flow makes it possible to obtain a very low content of NH ₃ dissolved at the bottom of the stripping column, making it possible to recover a third aqueous effluent which can be introduced into conventional wastewater treatment.

Step e) of NH3 separation is generally carried out at a pressure between 0.1 and less than 0.5 MPa, preferably between 0.05 and 0.2 MPa.

The flow rate of the inert gas flow is generally such that the ratio between the flow rate of the inert gas flow expressed in normal m ³ per hour (Nm ³ /h) and the flow rate of the load to be treated expressed in m ³ per hour at standard conditions (15 °C, 0.1 MPa) is between 50 and 600 Nm ³ /m ³ , preferably between 200 and 400 Nm ³ /m ³ . By normal m ³ we mean the quantity of gas in a volume of 1 m ³ at 0°C and 0.1 MPa.

According to a variant, in step e) a part of the gas phase at the top of the stripping column comprising NH3 and the inert gas flow (steam) is condensed and is preferably reinjected at least in part as liquid reflux in the upper part of the stripping column. Condensation is generally carried out by cooling to a temperature between 30 and 65°C, for example by cold water.

According to a variant, said gas phase containing NH3 can be at least partly recycled upstream of step a) and/or step b) and/or step g), and/or even upstream of the selective hydrogenation step aO), advantageously in stoichiometric quantities adapted to the formation of salts during separation/washing step c).

Step f) (optional) of splitting

The process according to the invention may comprise a step of fractionating all or part, preferably all, of the hydrocarbon effluent resulting from step c), to obtain at least one gaseous stream and at least two liquid hydrocarbon streams, said two liquid hydrocarbon streams being at least a first hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C (naphtha cut), in particular between 80 and 175°C, and a second hydrocarbon cut comprising compounds having a boiling point greater than 175°C (middle distillates cut).

Step f) makes it possible in particular to eliminate gases dissolved in the hydrocarbon liquid effluent, such as for example ammonia, hydrogen sulphide and light hydrocarbons having 1 to 4 carbon atoms.

The optional fractionation step f) is advantageously carried out at a pressure less than or equal to 3.0 MPa abs., preferably between 0.5 and 2.5 MPa abs.

According to one embodiment, step f) can be carried out in a section advantageously comprising at least one stripping column equipped with a reflux circuit comprising a reflux drum. Said stripping column is fed by the liquid hydrocarbon effluent from step c) and by a flow of water vapor. The liquid hydrocarbon effluent from step c) can possibly be heated before entering the stripping column. Thus, the lightest compounds are carried to the top of the column and into the reflux circuit comprising a reflux drum in which a gas/liquid separation takes place. The gas phase, which includes the light hydrocarbons, is withdrawn from the reflux drum in a gas stream. The hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C is advantageously withdrawn from the reflux flask. The hydrocarbon cut comprising compounds having a boiling point greater than 175°C is advantageously drawn off at the bottom of the stripping column.

According to other embodiments, fractionation step f) can use a stripping column followed by a distillation column or only a distillation column.

Depending on the destination or use of the cuts resulting from fractionation step f), those skilled in the art will adjust the cutting points in the stripping and/or distillation operations. For example, it may be necessary to adjust the naphtha cutting end point to 150, 175 or 200°C.

The first hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C and the second hydrocarbon cut comprising compounds having a boiling point greater than 175°C, optionally mixed, can be sent , in whole or in part, to a steam cracking unit, at the end of which Olefins can be (re)formed to participate in the formation of polymers. Preferably, only part of said cuts is sent to a steam cracking unit; at least a fraction of the remaining part is optionally recycled in at least one of the stages of the process and/or sent to a fuel storage unit, for example a naphtha storage unit, a diesel storage unit or a kerosene storage unit, from conventional oil charges.

According to a preferred mode, the first hydrocarbon cut comprising compounds having a boiling point less than or equal to 175 ° C, all or part, is sent to a steam cracking unit, while the second hydrocarbon cut comprising compounds having a boiling point boiling point greater than 175°C is sent to a hydrocracking step g) and/or sent to a fuel storage unit.

In a particular embodiment, the optional fractionation step f) can make it possible to obtain, in addition to a gas flow, a naphtha cut comprising compounds having a boiling point less than or equal to 175°C, preferably between 80 and 175°C, and, a middle distillates cut comprising compounds having a boiling point greater than 175°C and less than 385°C, and a hydrocarbon cut comprising compounds having a boiling point greater than or equal to 385°C, called heavy hydrocarbon cut. 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 oil feedstocks, it can still be recycled; the middle distillate cut can also be, in whole or in part, either sent to a steam cracking unit, or to a diesel storage unit from conventional oil feedstocks, or even be recycled; the heavy cut can for its part be sent, at least in part, to a steam cracking unit, or be sent to step g) of hydrocracking when it is present.

In another particular embodiment, the optional fractionation step f) can make it possible to obtain, in addition to a gas flow, a naphtha cut comprising compounds having a boiling point less than or equal to 175°C, preferably between 80 and 175°C, and a kerosene cut comprising compounds having a boiling point greater than 175°C and less than or equal to 280°C, a diesel cut comprising compounds having a boiling point greater than 280°C and less than 385°C and a hydrocarbon cut comprising compounds having a boiling point greater than or equal to 385°C, called heavy hydrocarbon cut. The naphtha cut, the kerosene cut and/or the diesel cut can (may) be, in whole or in part, either sent to a steam cracking unit, or respectively to a naphtha, kerosene or diesel pool from conventional oil feeds , or be recycled. The heavy cut can be sent, at least in part, to a steam cracking unit, or be sent to hydrocracking step g) when it is present.

In another particular embodiment, the naphtha cut comprising compounds having a boiling point less than or equal to 175°C from step f) is fractionated into a heavy naphtha cut comprising compounds having a boiling point between 80 and 175°C and a light naphtha cut comprising compounds having a boiling point below 80°C, at least part of said heavy naphtha cut being sent to an aromatic complex comprising at least one step of reforming the naphtha into to produce aromatic compounds. According to this embodiment, at least part of the light naphtha cut is sent to the steam cracking step h) described below.

The gaseous effluent(s) resulting from fractionation step f) may be subject to additional purification(s) and separation(s) in order to recover at least light hydrocarbons, in particular ethane, propane and butane, which can advantageously be sent separately or mixed into one or more ovens of steam cracking step h) so as to increase the overall yield of olefins.

(optional) step g) hydrocracking

According to a variant, the process of the invention may comprise a hydrocracking step g) carried out after step c) of separation with at least part of said hydrocarbon effluent from step c) or carried out after step f ) fractionation with at least part of the second hydrocarbon cut comprising compounds having a boiling point greater than 175°C.

Advantageously, step g) implements hydrocracking reactions well known to those skilled in the art, and more particularly makes it possible to convert heavy compounds, for example compounds having a boiling point greater than 175°C into compounds having a boiling point less than or equal to 175°C contained in the hydrocarbon effluent resulting from fractionation step f). Other reactions, such as hydrogenation of olefins, aromatics, hydrodemetallation, hydrodesulfurization, hydrodenitrogenation, etc. can continue.

Compounds with a boiling point above 175°C have a high BMCI and contain, compared to lighter compounds, more naphthenic, naphtheno-aromatic and aromatic compounds thus leading to a higher C/H ratio. This high ratio is a cause of coking in the steam cracker, thus requiring steam cracking ovens dedicated to this cut. When we wish to minimize the yield of these heavy compounds (middle distillate cut) and maximize the yield of light compounds (naphtha cut), these compounds can be transformed at least in part into light compounds by hydrocracking, a cut generally favored for a steam cracking unit.

Thus, the process of the invention may comprise a hydrocracking step g) implemented in a hydrocracking reaction section, implementing at least one fixed bed reactor having n catalytic beds, n being a greater integer or equal to 1, each comprising at least one hydrocracking catalyst, said hydrocracking reaction section being fed by at least part of said hydrocarbon effluent from step c) and/or by at least part of the second hydrocarbon cut comprising compounds having a boiling point greater than 175°C from step f) and a 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 volume velocity between 0.1 and 10.0 ^h'1 to obtain a first hydrocracked effluent.

Thus, said hydrocracking reaction section is advantageously carried out at an average temperature between 250 and 450°C, preferably between 320 and 430°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 an hourly volume velocity (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 g) is advantageously between 80 and 2000 Nm ³ of hydrogen per m ³ of fresh feed which feeds step a), and preferably between 200 and 1800 Nm ³ of hydrogen per m ³ of fresh charge which feeds stage a). The definitions of mean temperature (WABT), WH and hydrogen coverage correspond to those described above. Advantageously, said hydrocracking reaction section is carried out at a pressure equivalent to that used in the reaction section of step a) of hydrogenation or of step b) of hydrotreatment.

Advantageously, said step g) is implemented 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, catalyst(s) d hydrocracking. 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. The hydrocracked effluent can at least partly be recycled in step a) of hydrogenation and/or in step b) of hydrotreatment and/or in step c) of separation. Preferably, it is recycled in separation step c).

The hydrocracking step can be carried out in one (step g) or two steps (step g) and g')). When it is carried out in two stages, a separation is carried out of the effluent resulting from the first hydrocracking stage g) making it possible to obtain a hydrocarbon cut comprising compounds having a boiling point greater than 175°C (cut middle distillates), which is introduced into the second hydrocracking step g') comprising a second dedicated hydrocracking reaction section, different from the first hydrocracking reaction section g). This configuration is particularly suitable when you only want to produce a naphtha cut.

The second hydrocracking step g') is implemented in a hydrocracking reaction section, using 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 the first hydrocracked effluent from the first hydrocracking step g) and a gas stream comprising hydrogen, said hydrocracking reaction section being implemented 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 volume velocity between 0.1 and 10.0 h- ¹ , to obtain a second hydrocracked effluent. The preferred operating conditions and catalysts used in the second hydrocracking stage are those described for the first hydrocracking stage. The operating conditions and catalysts used in the two hydrocracking stages may be identical or different.

Said second hydrocracking step is preferably implemented in a hydrocracking reaction section comprising at least one, preferably between one and five, fixed bed reactor(s) having n catalytic beds, n being a greater integer 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, catalyst ( s) hydrocracking.

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 less than 160°C and preferably less than 150°C, greater than 15% by weight and even more preferably between 20 and 95% by weight. When the process is carried out in two hydrocracking stages, the conversion passes in the second stage is kept moderate in order to maximize the selectivity for compounds of the naphtha cut (having a boiling point less than or equal to 175°C, in particular between 80°C and less than or equal to 175°C). The conversion per pass is limited by the use of a high recycle rate on the second hydrocracking stage loop. This rate is defined as the ratio between the supply flow rate of step g') and the load flow rate of step a) or step b), preferably this ratio is between 0.2 and 4, preferably between 0.5 and 2.5.

The hydrocracked effluent from the second hydrocracking step g') can at least partly be recycled in the hydrogenation step a) and/or in the hydrotreatment step b and/or in the step c) of seperation. Preferably, it is recycled in separation step c).

The hydrocracking step(s) therefore does not necessarily make it possible to transform all the hydrocarbon compounds having a boiling point higher than 175°C (middle distillate cut) into hydrocarbon compounds having a lower boiling point. or equal to 175°C (naphtha cut). After the fractionation step f), there may therefore remain a more or less significant proportion of compounds having a boiling point greater than 175°C. To increase the conversion, at least part of this unconverted cut can be introduced into a second hydrocracking step g’). Another part can be purged. Depending on the operating conditions of the process, said purge may be between 0 and 10% by weight of the cut comprising compounds having a boiling point greater than 175°C relative to the incoming charge, and preferably between 0.5 % and 5% weight.

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

The hydrocracking catalyst(s) used in the hydrocracking step(s) are conventional hydrocracking catalysts known to those skilled in the art, of the bifunctional type combining an acid function with a function hydro-dehydrogenating agent and optionally at least one binder matrix. The acid function is provided by supports with a large surface area (150 to 800 m ² /g generally) presenting a superficial acidity, such as halogenated aluminas (chlorinated or fluorinated in particular), combinations of boron and aluminum oxides, amorphous silica-aluminas and zeolites. The hydro-dehydrogenating function is provided by at least one metal from group VI B of the periodic table and/or at least one metal from group VIII.

Preferably, the hydrocracking catalyst(s) comprise a hydrodehydrogenating function comprising at least one Group VIII metal chosen from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum, and preferably among cobalt and nickel. Preferably, said catalyst(s) also comprise at least one Group VI B metal chosen from chromium, molybdenum and tungsten, alone or as a mixture, and preferably from molybdenum and tungsten. Hydro-dehydrogenating functions of the NiMo, NiMoW, NiW type are preferred.

Preferably, the content of Group VIII metal in the hydrocracking catalyst(s) is advantageously between 0.5 and 15% by weight and preferably between 1 and 10% by weight, the percentages being expressed as a percentage by weight 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.

Preferably, the content of Group VI B metal in the hydrocracking catalyst(s) is advantageously between 5 and 35% by weight, and preferably between 10 and 30% by weight, the percentages being expressed as a percentage by weight of oxides relative to the total weight of the catalyst. When the metal is molybdenum or tungsten, the metal content is expressed as MoOs and WO3 respectively.

The hydrocracking catalyst(s) may also optionally comprise at least one promoter element deposited on the catalyst and chosen from the group formed by phosphorus, boron and silicon, optionally at least one element from group VI IA (chlorine , preferred fluorine), optionally at least one element from group VI I B (preferred manganese), and optionally at least one element from group VB (preferred niobium).

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

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

Preferably, the hydrocracking catalyst(s) also optionally comprise a zeolite chosen from Y zeolites, preferably from USY zeolites, alone or in combination, with other zeolites from beta zeolites, ZSM-12, IZM-2, ZSM-22, ZSM-23, SAPO-11, ZSM-48, ZBM-30, alone or in mixture. Preferably the zeolite is USY zeolite alone.

In the case where said catalyst comprises 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. A preferred catalyst comprises, and preferably consists of, at least one Group VI B metal and optionally at least one non-noble Group VIII metal, at least one promoter element, and preferably phosphorus, d at least one Y zeolite and at least one alumina binder.

An even more preferred catalyst comprises, and preferably consists of, nickel, molybdenum, phosphorus, a USY zeolite, and optionally also a beta zeolite, and alumina.

Another preferred catalyst includes, and preferably consists of, nickel, tungsten, alumina and silica-alumina.

Another preferred catalyst includes, and preferably consists of, nickel, tungsten, USY zeolite, alumina and silica-alumina.

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

In a variant, the hydrocracking catalyst used in the second hydrocracking step comprises a hydro-dehydrogenating function comprising at least one noble metal from group VIII chosen from palladium and platinum, alone or as a mixture. The content of noble metal from group VIII is advantageously between 0.01 and 5% by weight and preferably between 0.05 and 3% by weight, the percentages being expressed as a percentage by weight of oxides (PtO or PdO) relative to the weight. total catalyst.

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 “additive catalyst”. Generally, the organic compound is chosen from a compound comprising one or more chemical functions chosen from a carboxylic function, alcohol, thiol, thioether, sulfone, sulfoxide, ether, aldehyde, ketone, ester, carbonate, amine, nitrile, imide, oxime, urea and amide or even compounds including a furan cycle or even sugars.

The preparation of catalysts for the hydrogenation, hydrotreatment and hydrocracking stages is known and generally includes a step of impregnation of metals from group VIII and group VI B when present, and optionally phosphorus and/or boron on the support, followed by drying, then possibly calcination. In the case of additive catalyst, the preparation is generally done by simple drying without calcination after introduction of the organic compound. Here, calcination means heat treatment under a gas containing air or oxygen at a temperature greater than or equal to 200°C. Before their use in a step of the process, the catalysts are generally subjected to sulfurization in order to form the active species. The catalyst of step a) can also be a catalyst used in its reduced form, thus involving a reduction step in its preparation.

The gas stream comprising hydrogen, which feeds the selective hydrogenation, hydrogenation, hydrotreatment and hydrocracking reaction section may consist of additional hydrogen and/or recycled hydrogen originating in particular from step c) of separation. Preferably, an additional gas flow comprising hydrogen is advantageously introduced at the inlet of each reactor, in particular operating in series, and/or at the inlet of each catalytic bed from the second catalytic bed of the reaction section. These additional gas flows are also called cooling flows. They make it possible to control the temperature in the reactor in which the reactions carried out are generally very exothermic.

Said hydrocarbon effluent or said hydrocarbon stream(s) thus obtained by treatment according to the process of the invention of a plastic pyrolysis oil, has a composition compatible with the specifications of an input charge to a steam cracking unit. In particular, the composition of the hydrocarbon effluent or said hydrocarbon stream(s) is preferably such that:

- the total content of metallic elements is less than or equal to 10.0 ppm by weight, preferably less than or equal to 2.0 ppm by weight, preferably less than or equal to 1.0 ppm by weight and preferably less than or equal to 0, 5 ppm by weight, with: a content of silicon element (Si) 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/or an iron element content (Fe) less than or equal to 200 ppb by weight,

- the sulfur content is less than or equal to 100 ppm by weight, preferably less than or equal to 50 ppm by weight, and/or

- 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 and/or

- the asphaltene content is less than or equal to 5.0 ppm by weight, and/or

- the total chlorine element content is less than or equal to 10 ppm by weight, preferably less than 1.0 ppm by weight, and/or

- a mercury content less than or equal to 5 ppb by weight, preferably less than 3 ppb by weight, and/or

- 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 contents are given in relative weight concentrations, percentage (%) weight, part(s) per million (ppm) weight or part(s) per billion (ppb) weight, relative to the total weight of the flow considered.

The process according to the invention therefore makes it possible to treat plastic pyrolysis oils to obtain an effluent which can be injected, in whole or in part, into a steam cracking unit.

(optional) heavy metal adsorption step

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

The gaseous effluents may in particular be the gaseous effluent from step c) and/or the gas phase containing H2S from step d) and/or the gas phase containing NH3 from step e) and/or the gaseous effluent from step f) of fractionation.

The liquid effluents may in particular be the hydrocarbon effluent from step c) and/or the first and/or the second hydrocarbon cut from step f).

The optional adsorption step makes it possible to eliminate or reduce the quantity of metallic impurities, in particular the quantity of heavy metals such as arsenic, zinc, lead, and in particular mercury, possibly present in said effluents. gases and liquids. Metallic impurities, and particularly heavy metals, are present in the filler. Certain impurities, in particular based on mercury, can be transformed in one of the stages of the process according to the invention. Their transformed form is easier to trap. Their elimination or reduction may in particular be necessary when at least part of said gaseous and liquid effluents is intended to be sent, either directly, or after having undergone one or more optional additional steps such as the fractionation step f), in a step having specifications for severe metallic impurities, such as a steam cracking step.

Thus, an optional step of adsorption of a gaseous effluent and/or a hydrocarbon effluent resulting from the process according to the invention is advantageously carried out in particular when at least one of these effluents or the load respectively comprises more than 20 ppb by weight , in particular more than 15 ppb weight of metallic elements of heavy metals (As, Zn, Pb, Hg, etc.), and in particular when at least one of these effluents or the load respectively comprises more than 10 ppb weight of mercury , more particularly more than 15 ppb weight of mercury. Said optional adsorption step is advantageously carried out at a temperature between 20 and 250°C, preferably between 40 and 200°C, and at a pressure between 0.15 and 10.0 MPa abs, preferably between 0. 2 and 1.0 MPa abs.

Said optional adsorption step can be implemented by any adsorbent known to those skilled in the art making it possible to reduce the quantity of such contaminants.

According to a 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 which may be based on sulfur in elementary form. , or in the form of metal sulphide or metal oxide, or even in metallic form in elementary form.

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

The active phase is based on sulfur in elemental form, or in the form of metal sulphide or metal oxide, or even in metallic form in elemental form. Preferably, the active phase is in the form of a metal sulphide, in particular a sulphide of a metal from the group chosen from copper, molybdenum, tungsten, iron, nickel or cobalt.

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 very preferably between 3 and 20%.

Advantageously, 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%, preferably between 5 and 30%, very preferably between 5 and 20%.

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

Said adsorption section may comprise one or more adsorption columns. When the adsorption section comprises two adsorption columns, one mode of operation may be an operation called "swing", according to the established Anglo-Saxon term, in which one of the columns is online, that is to say in operation, while the other column is in reserve. Another mode of operation is to have at least two columns operating in series in swappable mode.

Preferably, said adsorption section comprises an adsorption column for the gaseous effluent(s) and an adsorption column for the liquid effluent(s).

Step h) of steam cracking (optional)

The hydrocarbon effluent from step c) of separation, or at least one of the two liquid hydrocarbon streams from step f) optional, can be sent in whole or in part to a step h) steam cracking.

Advantageously, the gaseous effluent(s) resulting from step c) of separation and/or f) of fractionation and containing ethane, propane and butane, can (can) in all or part also be sent to step h) of steam cracking.

Said steam cracking step h) is advantageously carried out in at least one pyrolysis oven at a temperature between 700 and 900°C, preferably between 750 and 850°C, and at a pressure between 0.05 and 0.3 MPa relative. The residence time of hydrocarbon compounds is generally less than or equal to 1.0 seconds (denoted s), preferably between 0.1 and 0.5 s. Advantageously, water vapor is introduced upstream of the optional steam cracking step h) and after the separation step c) (or fractionation f)). The quantity of water introduced, advantageously in the form of water vapor, is advantageously between 0.3 and 3.0 kg of water per kg of hydrocarbon compounds at the input of step h). Preferably, the optional step h) is carried out in several pyrolysis ovens in parallel so as to adapt the operating conditions to the different flows supplying step h) in particular from step f), and also to manage the processing times. decoking of the tubes. 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 can be dedicated to cracking the hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C.

The effluents from the different steam cracking furnaces are generally recombined before separation in order to constitute an effluent. It is understood that steam cracking step h) includes the steam cracking furnaces but also the sub-steps associated with steam cracking well known to those skilled in the art. These sub-stages may include heat exchangers, columns and catalytic reactors and recycling to the ovens. A column generally makes it possible to fractionate the effluent 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 optionally a fraction comprising pyrolysis oil. Columns make it possible to separate the different constituents of the light fractionation fraction in order to recover at least one cut rich in ethylene (C2 cut) and one cut rich in propylene (C3 cut) and possibly a cut rich in butenes (C4 cut). Catalytic reactors make it possible in particular to carry out hydrogenation of C2, C3 or even C4 cuts and pyrolysis gasoline. The saturated compounds, in particular the saturated compounds having 2 to 4 carbon atoms, are advantageously recycled to the steam cracking furnaces so as to increase the overall yields of olefins.

This steam cracking step h) makes it possible to obtain at least one effluent containing olefins comprising 2, 3 and/or 4 carbon atoms (that is to say C2, C3 and/or C4 olefins), at satisfactory contents, in particular greater than or equal to 30% by weight, in particular greater than or equal to 40% by weight, or even greater than or equal to 50% by weight of total olefins comprising 2, 3 and 4 carbon atoms relative to the weight of the effluent of steam cracking considered. Said C2, C3 and C4 olefins can then be advantageously used as polyolefin monomers.

According to a preferred embodiment of the invention, the process for treating a load comprising a plastic pyrolysis oil preferably comprises the sequence of steps, and preferably in the given order:

- b) hydrotreatment, c) separation/washing and d) separation of H2S with recycling of H2S in step b), and e) separation of NH3

- a) hydrogenation, b) hydrotreatment, c) separation/washing and d) separation of H ₂ S with recycling of H ₂ S in step a) and/or b), and e ) separation of NH ₃

- a) hydrogenation, b) hydrotreatment, c) separation/washing and d) separation of H ₂ S with recycling of H ₂ S in step a) and/or b), e) separation of NH ₃ and f) fractionation

- a) hydrogenation, b) hydrotreatment, c) separation/washing and d) separation of h S with recycling of H2S in step a) and/or b), e) separation of NH ₃ and f) fractionation and introduction of the hydrocarbon cut comprising compounds having a boiling point greater than 175°C in step g) of hydrocracking, the hydrocracked effluent being recycled in step c).

All embodiments may comprise and preferably consist of more than one pretreatment step aO). All embodiments may comprise and preferably consist of more than one steam cracking step h).

Analysis methods used

The analysis 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 those skilled in the art. They are specifically listed below for information purposes. Other methods deemed equivalent may also be used, in particular IP, EN or ISO equivalent methods:

Table 1

(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. 1, pp. 57-68 LIST OF FIGURES

The mention of the elements referenced in Figures 1 to 2 allows a better understanding of the invention, without it being limited to the particular embodiments illustrated in Figures 1 to 2. The different embodiments presented can be used alone or in combination with each other, without limitation of combination.

Figure 1 represents the diagram of a particular embodiment of the method of the present invention, comprising:

- a step a) of hydrogenation (optional) of a hydrocarbon feedstock 1 resulting from the pyrolysis of plastics in the presence of a hydrogen-rich gas 2 and possibly an amine provided by flow 3 and possibly an agent sulfurizing by flow 4 (particularly at the start of the cycle);

-a step b) hydrotreatment supplied by the hydrocarbon effluent 5 from step a) of hydrogenation if present and by a flow of a hydrogen-rich gas 6;

-a separation step c) supplied by the effluent 7 from the hydrotreatment step b) and in the presence of an aqueous solution 10 to obtain at least one gaseous effluent 11, a first aqueous effluent 12 containing I' H2S, HCl and NHs and a hydrocarbon effluent 13;

- a step d) of separating the H2S contained in the first aqueous effluent 12 preferably by stripping with a flow containing water vapor 19 making it possible to obtain a gas phase containing the H2S 20 and a second aqueous effluent 21, said gas phase containing H2S is at least partly recycled upstream of step a) and/or step b), preferably upstream of step a) when it is present. This recycling of the phase containing H ₂ S 20 makes it possible to maintain the catalysts of steps a) and/or b) in sulphide form and thus to reduce the supply of sulphurizing agent 4;

- a step e) of separation of the NH3 contained in the second aqueous effluent 21 preferably by stripping with a flow containing water vapor 19 making it possible to obtain a gas phase containing NH3 22 and a third aqueous effluent 23. Figure 2 represents the diagram of another particular embodiment of the process of the present invention which is based on the diagram of Figure 1. This diagram includes a step c) carried out in two steps, then a step f) of fractionation and an additional hydrocracking step g).

Hydrogenation step a) and hydrotreatment step b) are carried out as described in Figure 1. Separation step c), carried out in two steps, includes in particular:

- a step c1) of separation of the hydrotreated effluent 7 carried out at high pressure and high temperature (H H PS) to obtain at least one gaseous effluent 8, and a liquid effluent 9, part of which 9a can be recycled upstream of the step a) or upstream of step b) (not shown),

- a separation step c2) carried out at high pressure and low temperature (CH PS) and supplied by the gaseous effluent 8 and the other part of the liquid effluent 9b from step c1) and an aqueous solution 10 and making it possible to obtain at least one gaseous effluent 11 comprising hydrogen, an aqueous effluent 12 containing dissolved salts and dissolved H2S and NHs, and a hydrocarbon effluent 13;

Step d) of separation of H2S and step e) of separation of N H3 are carried out as described in Figure 1. Recycling of the phase containing H2S 20 is carried out in the same way. It can also at least partly be recycled in step g) of hydrocracking.

Optionally, a step f) of fractionating the hydrocarbon effluent 13 is carried out, making it possible to obtain at least one gaseous effluent 14, a first hydrocarbon cut 15 comprising compounds having a boiling point less than or equal to 175°C (cut naphtha) and a second hydrocarbon cut 16 comprising compounds having a boiling point greater than 175°C (middle distillates cut).

At the end of step f), a part of the first hydrocarbon cut 15 comprising compounds having a boiling point less than or equal to 175°C can be sent to a steam cracking process (not shown). Another part of the first hydrocarbon cut 15 can feed the hydrogenation step a) and/or the hydrotreatment step b) (recycling not shown).

In Figure 2, at least part of the second hydrocarbon cut 16 comprising compounds having a boiling point greater than 175°C from step f) feeds a hydrocracking step g) which is carried out in at least a fixed bed reactor comprising at least one hydrocracking catalyst and is supplied with hydrogen 17. The hydrocracked effluent 18 can be recycled between separation stages c1) and c2) or upstream of separation stage c) (not shown). Instead of injecting the amine stream 3 at the inlet of step a) of hydrogenation, it is possible to inject it at the inlet of step b) of hydrotreatment, at the inlet of step c) separation, at the entrance to step g) of hydrocracking when it is present, or even not to inject it, depending on the characteristics of the feed. Only the main steps, with the main flows, are represented in Figures 1 and 2, in order to allow a better understanding of the invention. It is of course understood that all the equipment necessary for operation is present (tanks, pumps, exchangers, ovens, columns, etc.), even if not shown. It is also understood that streams of hydrogen-rich gas (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. Means well known to those skilled in the art for purifying and recycling hydrogen can also be used.

EXAMPLES

Example 1 (in accordance with the invention)

Charge 1 treated in the process with a flow rate of 10,000 kg/h (10 T/h) is a plastic pyrolysis oil (that is to say comprising 100% weight of said plastic pyrolysis oil) having the characteristics shown in Table 2.

Table 2: load characteristics

(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. 1, pp. 57-68

Charge 1 is subjected to a hydrogenation step a) carried out in a fixed bed reactor and in the presence of hydrogen 2 and a NiMo type hydrogenation catalyst on alumina under the conditions indicated in Table 3.

Table 3: conditions of hydrogenation step a)

The conditions indicated in Table 3 correspond to cycle start conditions and the average temperature (WABT) is increased by 1°C per month in order to compensate for the catalytic deactivation.

At the end of step a) of hydrogenation, the conversion rates (= (initial concentration - final concentration) / initial concentration) observed are indicated in Table 4.

Table 4: species conversions during hydrogenation step a)

The effluent 5 from step a) of hydrogenation is subjected directly, without separation, to a step b) of hydrotreatment carried out in a fixed bed and in the presence of hydrogen and a NiMo type hydrotreatment catalyst. on alumina under the conditions presented in table 5. Table 5: conditions of hydrotreatment step b)

The conditions indicated in Table 5 correspond to cycle start conditions and the average temperature (WABT) is increased by 1°C per month in order to compensate for the catalytic deactivation.

The effluent 7 from step b) of hydrotreatment is subjected to a step c) of separation: a flow of water 10 is injected into the effluent from step b) of hydrotreatment; the mixture is then treated in an acid gas washing column and separator flasks to obtain a gas fraction and a liquid effluent. The yields of the different fractions obtained after separation are indicated in Table 6 (the yields correspond to the ratios of the mass quantities of the different products obtained relative to the charge mass upstream of step a), expressed as a percentage and denoted as % m /m).

Table 6: yields of the different products obtained after separation

All or part of the liquid fraction obtained can then be valorized in a steam cracking step with a view to forming olefins which can be polymerized with a view to forming recycled plastics.

The pyrolysis oil feed contains very little sulfur (170 ppm weight). This sulfur, which is in the form of sulfur molecules, is hydrogenated in the reaction section and is transformed into H ₂ S. This H ₂ S, in the form of partial pressure of H ₂ S (ppH2S) in the reactor, participates in the maintenance of the sulphide phase of NiMo catalysts on alumina. However, the ppH ₂ S obtained with this sulfur content in the feed (170 ppm by weight) is insufficient to maintain the catalysts in the sulphide phase throughout the cycle. This results in a rapid deactivation of the catalyst activity if nothing is done. It is therefore necessary to add H ₂ S to the reaction system in order to achieve a sufficient ppH ₂ S. This addition of H ₂ S can be done in the form of an injection at the inlet of the unit into the di-Methyl disulfide (DMDS) pyrolysis oil charge. DMDS decomposes easily upon contact on the catalyst in CH4 and H2S thus generating a PPH2S sufficient to maintain the catalysts in sulphide form. This way of doing things leads to a high consumption of DMDS which is detrimental to the economics of the process.

Another way which is the subject of the invention is to recover the H2S which is evacuated into the aqueous effluent using double stripping of this aqueous effluent and to re-inject this H ₂ S into the aqueous effluent. entry to the unit by dissolution in the pyrolysis oil charge.

The injection of DMDS and/or the recycling of H ₂ S at the inlet of the unit can be used both to maintain a sufficient ppH ₂ S in the reaction system but also can be used to neutralize all the NH ₃ resulting hydrogenation of nitrogen molecules. Indeed, the H ₂ S reacts with the NHs to form ammonium sulphides which will be almost completely washed and transferred into the aqueous effluent (flow 12) allowing the gas flow to be released at the head outlet of the column. stabilization of the presence of ammonia (flow 11). This gas flow freed from the presence of ammonia can therefore be sent directly to the steam cracker in order to maximize olefin production.

The advantage of recycling a stream of h^S compared to injecting DMDS, whether in the context of maintaining a sufficient pph^S or in the case of delivering a gas stream stripped of its ammonia, is therefore to save the quantity of DMDS throughout the cycle.

Table 7 shows four operating cases.

Case 1: Stripper of simple acidic water and injection of DMDS only to maintain sufficient PPH2S to keep the catalysts in the sulphide phase

Case 2: Double acidic water stripper to recycle at the entrance to the unit a majority flow of H ₂ S coming from the head of the first stripping column only to maintain a sufficient ppH ₂ S to keep the catalysts in phase sulfide. This case is consistent with the invention.

Case 3: Stripper of simple acidic water and injection of DMDS to maintain a sufficient ppH ₂ S to keep the catalysts in the sulphide phase and also to neutralize all the NH3 and deliver a gas flow without NH3

Case 4: Double acid water stripper to recycle at the entrance to the unit a majority flow of H2S coming from the head of the first stripping column to maintain a sufficient PPH2S to keep the catalysts in the sulphide phase and also in order to neutralize all NH3 and deliver a gas flow without NH3. This case is consistent with the invention.

It can be noted that the invention makes it possible to save 19 kg/h of DMDS when it comes to maintaining a minimum ppH2S to keep the catalysts in sulphide form. It can also be noted that the invention allows an even greater gain, namely 65 kg/h (75 - 10 = 65 kg/h) of DMDS saved to deliver a gas flow without ammonia.

Table 7: Operating cases

Operating conditions of the Simple Stripping column (H2S & NH3)

Claims

1. Process for treating a feedstock comprising a plastics pyrolysis oil, comprising: a) optionally a hydrogenation step implemented in a hydrogenation reaction section, using 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 supplied at least by said feed and a gas flow comprising hydrogen, said reaction section hydrogenation 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 volume velocity between 0.1 and 10.0 h' ¹ , to obtain a hydrogenated effluent, b) a hydrotreatment step implemented in a hydrotreatment reaction section comprising at least one hydrotreatment catalyst, said hydrotreatment reaction section being supplied at least by the feed or said hydrogenated effluent from step a) and a gas flow comprising hydrogen, said hydrotreatment 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 volume velocity between 0.1 and 10.0 h' ¹ , to obtain a hydrotreated effluent, c) a separation step, supplied by the hydrotreated effluent from step b) and possibly by the hydrocracked effluent from step g) and an aqueous solution to obtain at least one gaseous effluent, a first aqueous effluent and a hydrocarbon effluent, d) a step of separating the H2S contained in the first aqueous effluent to obtain a gas phase containing H ₂ S and a second aqueous effluent, said gas phase containing H ₂ S being optionally at least partly recycled upstream of step a) and/or step b) and/or step g ), e) a step of separating the NH ₃ contained in the second aqueous effluent to obtain a gas phase containing NH ₃ and a third aqueous effluent, said gas phase containing NH ₃ being optionally at least partly recycled upstream of the step a) and/or step b) and/or step g), f) optionally a step of fractionating all or part of the hydrocarbon effluent from step c), to obtain at least one gaseous effluent and at least a first hydrocarbon cut comprising compounds having a boiling point less than or equal to 175°C and a second hydrocarbon cut comprising compounds having a boiling point greater than 175°C, g) optionally a hydrocracking step implemented in a hydrocracking 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 hydrocracking catalyst, said hydrocracking reaction section being fed by at least part of said hydrocarbon effluent from step c) and/or by at least part of the second hydrocarbon cut comprising compounds having a boiling point greater than 175°C from step f) and a gas flow 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 volume velocity between 0.1 and 10.0 ^h'1 to obtain a first hydrocracked effluent.

2. Method according to the preceding claim in which said gas phase containing H2S resulting from step d) is at least partly recycled upstream of step a) and/or step b) and/or step g).

3. Method according to one of the preceding claims comprising the hydrogenation step a).

4. Method according to one of the preceding claims comprising the fractionation step f).

5. Process according to one of the preceding claims comprising the hydrocracking step g).

6. Method according to one of the preceding claims in which step d) of separating the H2S contained in the first aqueous effluent is carried out by stripping said effluent with a flow containing water vapor at a pressure between 0.5 and 1 MPa and a temperature between 80 and 150°C.

7. Method according to one of the preceding claims in which step e) of separation of the NH ₃ contained in the second aqueous effluent is carried out by stripping said effluent with a flow containing water vapor at a pressure comprised between 0.1 and 0.5 MPa and a temperature between 80 and 150°C.

8. Method according to one of the preceding claims in which separation step c) comprises the following steps: c1) a separation step, supplied by the hydrotreated effluent from step b), said step being carried out at a temperature between 200 and 450°C and at a pressure substantially identical to the pressure of step b) to obtain at least one gaseous effluent and one liquid effluent, part of which is optionally recycled upstream of step a) and /or step b), c2) a separation step, supplied by the gaseous effluent from step c1) and another part of the liquid effluent from step c1) and an aqueous solution, said step being carried out at a temperature between 20 and less than 200°C, and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent, a first aqueous effluent and a hydrocarbon effluent,

9. Method according to one of the preceding claims, comprising at least one step aO) of pretreatment of the load comprising a plastic pyrolysis oil, optionally mixed with the hydrocarbon effluent from step c), said step of pretreatment being implemented upstream of step a) and/or upstream of step b) and comprises a filtration step and/or a centrifugation step and/or an electrostatic separation step and/or a step washing using an aqueous solution and/or an adsorption step and/or a selective hydrogenation step.

10. Method according to one of the preceding claims, in which the hydrocarbon effluent resulting from step c) of separation, or at least one of the two liquid hydrocarbon cuts resulting from step f ), is entirely or partly sent to a steam cracking step h) carried out in at least one pyrolysis oven at a temperature between 700 and 900°C and at a pressure between 0.05 and 0.3 MPa relative.

11. Method according to one of the preceding claims in which said gas phase containing NH3 from step e) is at least partly recycled upstream of step a) and/or step b) and/or l step g).

12. Method according to one of the preceding claims in which a flow containing a nitrogen compound and/or a sulfur compound is injected upstream of step a) and/or upstream of step b).

13. Method according to one of the preceding claims in which said hydrogenation catalyst comprises a support chosen from alumina, silica, silica-aluminas, magnesia, clays and their mixtures and a hydro-dehydrogenating function comprising either at least one element from group VIII and at least one element from group VIB, i.e. at least one element from group VIII.

14. Method according to one of the preceding claims in which said hydrotreatment catalyst comprises a support chosen from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and their mixtures, and a function hydro-dehydrogenating agent comprising at least one element from group VIII and/or at least one element from group VIB.

15. Method according to one of the preceding claims, which further comprises a second hydrocracking step g') implemented in a hydrocracking reaction section, using at least a 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 part of the first hydrocracked effluent resulting from the first hydrocracking step g) and a gas flow 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 volume velocity between 0.1 and 10.0 h- ¹ , to obtain a second hydrocracked effluent.

16. Method according to one of the preceding claims, in which said hydrocracking catalyst comprises a support chosen from halogenated aluminas, combinations of boron and aluminum oxides, amorphous silica-aluminas and zeolites and a function hydro-dehydrogenating agent comprising at least one metal from group VI B chosen from chromium, molybdenum and tungsten, alone or in a mixture, and/or at least one metal from group VIII chosen from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum.

17. Product obtained by the process according to one of claims 1 to 16.

18. Product according to claim 17, which comprises, relative to the total weight of the product:

- a silicon element content less than or equal to 5.0 ppm by weight, and/or

- a sulfur content less than or equal to 100 ppm by weight, and/or

- a nitrogen content less than or equal to 100 ppm by weight, and/or

- a chlorine element content less than or equal to 10 ppm by weight, and/or

- a mercury content less than or equal to 5 ppb by weight.