FR3144155A1 - METHOD FOR TREATMENT OF PYROLYSIS OILS OF PLASTICS AND/OR TIRES INCLUDING THE ELIMINATION OF HALIDES PRIOR TO A HYDROTREATMENT STEP - Google Patents

Abstract

Process for treating a plastics and/or tire pyrolysis oil comprising: a) the hydrogenation of said feed to obtain a hydrogenated effluent, b) a separation of the hydrogenated effluent carried out at a temperature higher than the temperature of precipitation of ammonium halides and at high pressure to obtain a gaseous effluent and a liquid effluent, c) the hydrotreatment of said liquid effluent to obtain a hydrotreated effluent, d) a separation of the hydrotreated effluent carried out at a temperature higher than the precipitation temperature of the ammonium halides and at high pressure to obtain a gaseous effluent and a liquid effluent, e) a separation carried out at a temperature lower than the precipitation temperature of the ammonium halides and high pressure, supplied by the gaseous effluents of steps b) and d) and the liquid effluent from step d) and an aqueous solution, to obtain a gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

Description

METHOD FOR TREATING PLASTICS AND/OR TIRES PYROLYSIS OILS INCLUDING THE ELIMINATION OF HALIDES PRIOR TO A HYDROTREATMENT STEP

The present invention relates to a process for treating a pyrolysis oil of plastics and/or tires in order to obtain a hydrocarbon effluent which can be valorized in a gasoline, jet or diesel fuel storage unit or as a charge for a steam cracking unit. More particularly, the present invention relates to a process for treating a charge resulting from the pyrolysis of plastic waste or tires in order to eliminate at least partly impurities, and in particular halides, between two catalytic stages using hydrogen .

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 may 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.

As for tires, they are mainly made up of gums for their elastic property (mixture of elastomers such as crosslinked natural and synthetic rubbers, with added additives such as silica, resin, sulfur, zinc oxide, carbon black, etc.) and textile and metal fibers for their reinforcing property.

Plastics from collection and sorting channels or recycled tires can undergo a pyrolysis stage in order to obtain, among other things, pyrolysis oils. These oils generally include a lot of impurities, in particular halogenated compounds, in particular compounds based on chlorine, but also diolefins, olefins, metals, in particular iron, silicon, or even heteroelements such as sulfur, oxygen and nitrogen, and insolubles.

These plastic and/or tire pyrolysis oils are generally burned to generate electricity and/or used as fuel in industrial or district heating boilers.

Another way of valorizing plastic and/or tire pyrolysis oils is the use of these pyrolysis oils as feedstock for a steam cracking unit in order to (re)create olefins, the latter being constituent monomers. of certain polymers.

Plastic and/or tire pyrolysis oils can also be used as gasoline, jet or diesel fuels.

However, plastic and/or tire pyrolysis oils have impurities at often high levels and are incompatible with direct storage in a fuel storage unit or with steam cracking units or units located downstream of the steam cracking units. , in particular polymerization processes and selective hydrogenation processes. These impurities can generate operability problems and in particular corrosion problems (in particular due to the presence of chlorine), coking or catalytic deactivation, or even incompatibility problems 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.

One way to eliminate these impurities contained in plastic and/or tire pyrolysis oils is to carry out hydrotreatment (HDT) in the presence of catalysts.

The most problematic impurities contained in plastic and/or tire oils are often halides and more particularly chlorine. Indeed, chlorine is generally the limiting contaminant for treating pyrolysis oils in a refinery's existing units. Chlorine, even at low levels (eg <5 ppm weight), is responsible for corrosion (in the form of HCl) which can occur in existing units whose metallurgy is generally not designed to resist chlorine levels. even weak. Another problem linked to the presence of halides in pyrolysis oils, and in particular chlorine, is the formation of ammonium chloride salts which form by reaction between chloride ions, released by hydrodechlorination in the form of HCl and ammonium ions, generated by the hydrogenation of nitrogen compounds (hydrodenitrogenation) in the form of NH ₃ during a hydrogenation and/or hydrotreatment step. It is known that these ammonium chloride salts precipitate at a relatively low temperature (eg less than 280°C) which creates blockage problems particularly in transfer lines and/or in sections of a process in process. downstream of hydrogenation/hydrotreatment.

Processes for eliminating chlorine by hydrotreatment in a pyrolysis oil are for example known from documents: WO20020769, WO20016400, WO20239729, WO21105326 or even WO16142809.

Document WO16142809 describes in particular a process including a pyrolysis step, a hydrotreatment step, a gas separation step (C1 to C4 containing H ₂ S and HCl) and a C5+ liquid, a dechlorination step of the C5+ liquid by adsorption or a second hydrotreatment step, possibly another step of separating the gases containing HCl, then a steam cracking step.

It is also known to eliminate the HCl formed during hydrotreatment by washing with water after a hydrotreatment step. Such washes are described in WO17083018, WO20254634 or even WO22101333. In none of these documents is washing before a hydrotreatment step described, nor is the problem of deactivation of the hydrotreatment catalyst of a second hydrotreatment step downstream of a first hydrotreatment step by a deposition of chlorinated compounds.

Unpublished patent application FR 21/12,908 describes a process for treating a plastic pyrolysis oil, comprising:

a) a step of hydrogenating the feed mixed with at least part of a liquid effluent from separation step c) and hydrogen,

b) a step of hydrotreating the effluent from step a) in the presence of hydrogen,

c) a separation step, supplied by the 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 a first gaseous effluent, and a liquid effluent, part of which is recycled upstream of step a),

d) a separation step, supplied by the first gaseous effluent and another part of the liquid effluent from step c) 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 c), to obtain at least a second gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

One of the objectives of application FR 21/12.908 is to eliminate, by the combination of step c) of hot separation followed by step d) of cold separation/washing, chlorine in the form of chloride salts ammonium. Chloride ions, released by the hydrogenation of chlorinated compounds in the form of HCl during steps a) and b) (hydrodechlorination) and ammonia generated by the hydrogenation of nitrogen compounds in the form of NH ₃ during step b) in particular (hydrodenitrogenation), largely leaves in the gaseous effluent thanks to the hot separation of step c). In fact, the high temperature of this separation step c) avoids the precipitation of ammonium chloride salt which is formed by reaction between the chloride ions and the ammonium ions. The separation at a lower temperature in step d) of the gaseous effluent and part of the liquid effluent causes these ammonium chloride salts to precipitate. Washing with water in this step d) makes it possible to dissolve these salts in the aqueous effluent. We thus obtain a hydrocarbon effluent free of chlorine.

This same principle of removing chlorine by hot stripping to keep the ammonia and halides in the gas phase until a quantity of water is available to collect the ammonium halides in solution is also described in WO22101333. In this document, this separation occurs, as in application FR 21/12.908, after two HDT stages with a view to sending the purified oil to a steam cracking stage, the objective being to avoid the precipitation of the halides of solid ammonium on internal surfaces of equipment and corrosion.

The present invention provides an improvement of the process according to FR 21/12.908 by proposing the elimination of chlorine not after step b) of hydrotreatment, but by a combination of a hot separation before step b) of hydrotreatment coupled with hot separation and cold separation with washing after hydrotreatment.

The applicant noticed that the chlorine contained in the feed is almost completely converted in the hydrogenation stage. She also observed that by sending the effluent from the hydrogenation stage directly into the hydrotreatment stage, the chlorine (in the form of HCl) contained in the effluent from the hydrogenation stage is deposited on the hydrotreatment catalyst thus leading to a drop in activity and a reduction in the duration of the catalytic cycle. In order to avoid clogging of the hydrotreatment catalyst and a drop in activity, the present invention proposes to eliminate the chlorine before the hydrotreatment step.

Objectives and Summary of the Invention

More specifically, the invention relates to a process for treating a load comprising a pyrolysis oil for plastics and/or tires comprising halogenated compounds, said process comprising:

a) 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 catalyst hydrogenation, said hydrogenation reaction section being supplied at least by said feed possibly mixed with at least part of the second liquid effluent from separation step d) and a first gas stream comprising hydrogen, said hydrogenation reaction section being carried out at an average temperature between 140 and 400°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volume velocity between 0.1 and 10.0 h ^-1 , to obtain a hydrogenated effluent,

b) a separation step, supplied by the hydrogenated effluent from step a), said step being carried out at a temperature higher than the precipitation temperature of the ammonium halides and at a pressure substantially identical to the pressure of the step a) to obtain at least a first gaseous effluent and a first liquid effluent,

c) a hydrotreatment step 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 catalyst hydrotreatment, said hydrotreatment reaction section being supplied at least by said first liquid effluent from step b), and a second gas stream comprising hydrogen, said hydrotreatment reaction section being implemented at a 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,

d) a separation step, supplied by the hydrotreated effluent from step c), said step being carried out at a temperature higher than the precipitation temperature of the ammonium halides and at a pressure substantially identical to the pressure of the step c) to obtain at least a second gaseous effluent and a second liquid effluent, part of which is optionally recycled upstream of step a),

e) a separation step, supplied by the first and second gaseous effluent and at least part of the second liquid effluent from step d) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature ammonium halides and at a pressure substantially identical to or lower than the pressure of step d), to obtain at least a third gaseous effluent, an aqueous effluent and a hydrocarbon effluent,

f) optionally a step of fractionating all or part of the hydrocarbon effluent from step e), to obtain at least a fourth gaseous effluent, a naphtha cut and at least one middle distillate cut,

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 hydrotreated effluent from step c) and/or by at least part of the middle distillate cut from step f) and a third gas stream comprising hydrogen, said hydrocracking reaction section being carried out at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly volume velocity between 0.1 and 10.0 h ^-1 to obtain a first hydrocracked effluent.

The major advantage of the present invention is to carry out the separation of the halogenated compounds contained in the hydrogenated effluent resulting from step a) before step c) of hydrotreatment, thus making it possible to protect the hydrotreatment catalyst while avoiding precipitation of solid ammonium halides on internal surfaces of equipment and corrosion.

The halogenated compounds contained in the feed are almost completely converted in step a) of hydrogenation into hydrogen halides (HCl for example), as well as part of the nitrogen compounds into gaseous ammonia NH ₃ . Hot separation (step b), that is to say at a temperature higher than the precipitation temperature of the ammonium halides, makes it possible to keep the ammonia and the halides in the gas phase and to obtain an essentially liquid effluent. freed from halogenated compounds which is sent to the hydrotreatment step in order to remove the remaining impurities without observing the deposition of halogenated compounds on the catalyst.

The hydrotreated effluent is then subjected again to a hot separation step (step d)), that is to say at a temperature higher than the precipitation temperature of the ammonium halides, which makes it possible to keep the contaminants (NH ₃ from HDN hydrodenitrogenation, H ₂ S from HDS hydrodesulfurization, etc.) in the gas phase and to obtain a liquid hydrocarbon effluent free of these contaminants.

The gaseous effluents resulting from the first and second hot separation (steps b) and d)) are then subjected to cold separation. Cold separation (step e), that is to say at a temperature lower than the precipitation temperature of the ammonium halides and in the presence of an aqueous solution, causes these ammonium halide salts to precipitate while by dissolving them in water.

The fact of eliminating halogenated compounds before the hydrotreatment step makes it possible to operate the hydrotreatment step c) at a lower average temperature than without eliminating the halogenated compounds beforehand, which presents an energy saving .

In addition, an increase in the cycle time of the hydrotreatment section is observed. By cycle duration we mean the duration of use of the catalyst without replacement of the catalyst being necessary.

During the first separation step b) at high temperature, almost all of the light fraction of the hydrogenated effluent (naphtha cut) leaves as the first gaseous effluent towards the cold separation/washing step e) while in the liquid phase we mainly have the heavy fraction of the feed (middle distillate cut). This last fraction is introduced as the first liquid effluent from step b) into hydrotreatment step c) to be hydrotreated. The first gaseous effluent from step b) is introduced into cold separation step e), and therefore bypasses hydrotreatment step c). By the process according to the invention, the first gaseous effluent from step b) is therefore not hydrotreated. The contaminants which are conventionally eliminated during hydrotreatment (nitrogen, sulfur, etc.) can therefore still be contained in this first gaseous effluent. Mixing the gaseous effluents, the first gaseous effluent from step b) (non-hydrotreated) and the second gaseous effluent from step d) (hydrotreated) with the second liquid effluent from step d) allows to achieve by dilution the specifications required for the use of pyrolysis oil in upstream units, for example in a steam cracking unit. The process according to the invention is therefore particularly suitable for “heavy” pyrolysis oil loads, that is to say comprising mainly a heavy fraction (medium distillate cut).

The fact that the gaseous effluent from step b) bypasses hydrotreatment step c) makes it possible to reduce the size of the reactor of the hydrotreatment step and therefore save on the need for catalyst and equipment.

Carrying out step e) of cold separation on the gaseous effluent(s) and liquid effluent(s) mixture (and not only on the gaseous effluent(s)) has the advantage of extracting not only the contaminants of the gaseous effluent(s), but also the hydrogen halides (HCl) dissolved in the hydrotreated liquid effluent.

Another major advantage of the present invention is to carry out the separation of the halogenated compounds of step b) at high pressure and in particular at a pressure substantially identical to the pressure of step a) of hydrogenation, which is also substantially identical to the pressure of hydrotreatment step c). The three steps a), b) and c) are thus carried out at the same pressure. The load therefore does not need to be depressurized between stages a) and c) which represents an energy saving.

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

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 the pyrolysis oil.

The process of the invention thus makes it possible to obtain a hydrocarbon effluent from a pyrolysis oil freed at least in part from the impurities of the starting pyrolysis oil, thus limiting operability problems, such as corrosion problems. , coking or catalytic deactivation, which these impurities can generate, in particular in the steam cracking units and/or in the units located downstream of the steam cracking units, in particular the 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, the process comprises the fractionation step f).

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

According to a variant, the process comprises a separation step b') carried out between steps b) and c), supplied with at least part of the first liquid effluent from step b) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of the ammonium halides and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

According to a variant, step b) or step d) is carried out at a temperature between 200 and 450°C and step e) or step b') is carried out at a temperature greater than or equal to 20 °C and less than 200°C.

According to one variant, at least part of the second liquid effluent obtained in step d) is recycled upstream of step a).

According to one variant, at least part of the hydrocarbon effluent from step e) is recycled upstream of step a) and/or upstream of step c).

According to a variant, in step a) the hydrogen coverage is between 250 and 800 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ).

According to one variant, the method comprises at least one step a0) of pretreatment of the pyrolysis oil fraction of plastics and/or tires, said pretreatment step being implemented upstream of step a) and comprises a adsorption step and/or a filtration step and/or a centrifugation step and/or a decantation step and/or an electrostatic separation step and/or a washing step using a solution aqueous and/or a gas stripping step.

According to a variant, the hydrocarbon effluent from step e) of separation, or at least one of the cuts resulting from step f), is sent in whole or in part to a step h) of steam cracking 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 one variant, the reaction section of step a) uses at least two reactors operating in switchable mode.

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

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

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 supplied 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 of the group VIB 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,
- a silicon element content less than or equal to 5.0 ppm by weight,
- a sulfur content less than or equal to 500 ppm by weight,
- a nitrogen content less than or equal to 100 ppm by weight,
- a chlorine element content less than or equal to 10 ppm by weight
- 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.

In this description, the term “comprising” is synonymous with (means the same as) “include” and “contain”, and is inclusive or open and does not exclude other elements that are not mentioned. It is understood that the term “understand” includes the exclusive and closed term “consist”.

In this description, the expression “between… and…” means that the limit values of the interval are included in the range of values described, unless otherwise specified.

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 rest of the text, “pyrolysis oil” means an oil resulting from the pyrolysis of plastics and/or tires, unless otherwise indicated.

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

The metal content is measured by X-ray fluorescence.

detailed description Load

According to the invention, a “plastic pyrolysis oil” or “tire 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 production lines. collection and sorting, or 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 in paraffins, up to 90% by weight in naphthenes, up to 90% by weight in olefins and up to 90% by weight in olefins. % weight of aromatics, it being understood that the sum of paraffins, naphthenes, olefins and aromatics is 100% weight of hydrocarbon compounds.

The oil may include diolefins. The diolefin content is commonly determined indirectly as the maleic anhydride index (or MAV for Maleic Anhydrid Value according to Anglo-Saxon terminology). The method is based on the Diels-Alder addition reaction between conjugated diolefins and maleic anhydride. The method for determining MAV is described in C. López-García et al., Near Infrared Monitoring of Low Conjugated Diolefins Content in Hydrotreated FCC Gasoline Streams, Oil & Gas Science and Technology – Rev. IFP, Vol. 62 (2007), No. 1, pp. 57-68. The MAV is expressed in mg of maleic anhydride reacted with 1 g of sample (mg/g). The MAV varies between 5 and 100 mg/g in pyrolysis oils.

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 pyrolysis oil can include, 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 pyrolysis oil at high levels, for example up to 600 ppm by weight or even 700 ppm by weight or even 1000 ppm by weight and even 5000 ppm by weight of halogen elements (in particular chlorine but also bromine , fluorine, iodine) provided by halogenated compounds, and generally between 1 and 1000 ppm by weight or between 1 and 700 ppm by weight or even between 1 and 600 ppm by weight of halogen elements. The pyrolysis oil can contain up to 600 ppm by weight or even 700 ppm by weight or even 1000 ppm by weight and even 5000 ppm by weight of chlorine element provided by chlorinated compounds, and generally between 1 and 1000 ppm by weight or between 1 and 700 ppm weight or between 1 and 600 ppm weight of chlorine elements.

The oil can include up to 200 ppm by weight, or even 1500 ppm by weight of metallic or semi-metallic elements, and generally between 1 and 200 ppm by weight or between 1 and 1500 ppm by weight of metallic or semi-metallic elements. Alkali metals, alkaline 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 include silicon, iron or both elements. The pyrolysis oil may in particular comprise up to 200 ppm by weight or even 1000 ppm by weight of silicon, and generally between 1 and 200 ppm by weight or between 1 and 1000 ppm by weight or even between 1 and 500 ppm by weight of silicon. The pyrolysis oil may in particular comprise up to 50 ppm by weight or even 100 ppm by weight of iron, and generally between 1 and 50 ppm by weight or between 1 and 100 ppm by weight of iron. Pyrolysis oil may also include phosphorus, sodium, calcium, potassium and magnesium.

The 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 less than 40,000 ppm weight of heteroelements and preferably less than 15,500 ppm weight of heteroelements, and generally between 1 and 40,000 ppm weight or between 1 and 15,500 ppm weight of heteroelements. The sulfur compounds are generally present in a content of less than 15,000 ppm by weight and preferably less than 10,000 ppm by weight, and generally between 1 and 15,000 ppm by weight or between 1 and 10,000 ppm by weight of sulfur compounds.

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, and generally between 1 and 15,000 ppm by weight or between 1 and 10,000 ppm by weight of oxygenated compounds.

The nitrogen compounds are generally present in a content of less than 10,000 ppm by weight and preferably less than 5,000 ppm by weight, and generally between 1 and 10,000 ppm by weight or between 1 and 5,000 ppm by weight of nitrogen compounds.

The 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 or arsenic , and generally between 1 and 200 ppb by weight or between 1 and 100 ppb by weight of heavy metals.

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

The feed for the process according to the invention may include, in addition to the 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 the feed.

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. Oils/fats of plant and/or animal origin contain triglycerides and/or free fatty acids and/or esters. 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, jatropha (jatropha), copra. , castor oil, 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. Algae or fish oils are also relevant. 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 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 hydrothermal liquefaction 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).

(a0) Pre-processing step (optional)

Said feed comprising a pyrolysis oil can advantageously be pretreated in an optional pretreatment step a0), prior to hydrogenation step a), to obtain a pretreated feed which feeds step a).

This optional pretreatment step a0) 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 plastic pyrolysis oil. Thus, an optional step a0) of pretreatment of the charge comprising a plastics pyrolysis oil is advantageously carried out in particular when said charge 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 a0) of pretreatment of the load comprising a plastic pyrolysis oil is advantageously carried out in particular when said load 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 a0) 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 an adsorption step and/or a filtration step and/or a centrifugation step and/or a decantation step and/or an electrostatic separation step and/or a washing step with using an aqueous solution and/or a gas stripping step.

The optional pretreatment step a0) 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, of preferably between 0.2 and 7.0 Mpa abs.

According to a variant, said optional pretreatment step a0) is implemented in an adsorption section operated in the presence of at least one adsorbent. The adsorbent can be chosen from a zeolite, activated carbon, clay, silica or alumina.

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 a0) 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, a mode of operation 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 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 a0) 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 a0) is implemented by filtration. The filtration step removes inorganic solids, sediments and/or fines contained in the oil, including metals, metal oxides and metal chlorides. A filter is generally used whose size (for example the diameter or equivalent diameter) of the pores is less than 25 µm, preferably less than or equal to 10 µm, even more preferably less than or equal to 5 µm. According to another variant, a filter can be used whose pore size is less than 25 µm but greater than 5 µm. A series of filters with different pore sizes can also be used, including a series of filters with decreasing pore sizes in the direction of oil flow. 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 heptane insoluble content should be reduced to less than 0.5% by weight, preferably less than 0.1%.

According to a particular embodiment, the pretreatment step a0) by filtration comprises at least one filter whose pore size is less than 10 microns, and preferably greater than 5 µm, possibly followed by a filtration system whose pore size is less than 2 µm and preferably less than 1 µm.

According to another particular embodiment, the pretreatment step a0) by filtration comprises at least one filter whose pore size is less than 10 µm, and preferably greater than 5 µm, followed by an electrostatic precipitation system.

According to another particular embodiment, the pretreatment step a0) by filtration comprises at least one filter whose pore size is less than 10 µm, and preferably greater than 5 µm, followed by a system of filter(s). ) using filter aids such as sand or diatomaceous earth.

According to another variant, said optional pretreatment step a0) is carried out by centrifugation. According to another variant, the pretreatment step a0) comprises centrifugation and filtration.

According to another variant, said optional pretreatment step a0) is implemented by decantation. According to another variant, the pretreatment step a0) comprises decantation and filtration.

According to another variant, said optional pretreatment step a0) is implemented by gas stripping, thus reducing the oxygen content in the oil. Gas extraction can remove oxygen (O ₂ ) which can be dissolved in the feed thus reducing the likelihood of free radical formation leading to polymerization in downstream stages. The process generally involves contacting the oil with an extracting gas (e.g. H ₂ , N ₂ or a mixture thereof), thereby transferring at least some of the dissolved oxygen from the oil to the stripping gas, followed by the separation of the stripping gas from the oil. The volume of extraction gas relative to the volume of oil (both volumes measured under gas extraction conditions) is generally greater than 1, and preferably at least 3. In particular embodiments, the extraction gas may contain at least 60% (molar percentage) H ₂ . Any dissolved H ₂ remaining in the feed after the gas stripping step is not a problem, considering the downstream hydrodemetallation/hydroprocessing. Preferably, the gas extraction step is completed before any (pre)heating of the feed, in order to minimize potential fouling.

Said optional pretreatment step a0) generally comprises one or more, preferably several treatments described above. It may in particular comprise a sequence of a washing step using an aqueous solution and/or an adsorption step, followed by a gas stripping step, followed by a filtration step and/or d a centrifugation step. All these steps are preferably carried out before any (pre)heating of the load.

Said optional pretreatment step a0) can also optionally be supplied with at least part of the liquid effluent from step d) of the process and/or part of at least one of the cuts resulting from the step e), mixed or separately from the charge comprising a pyrolysis oil. Recycling at least part of the liquid effluent from step d) makes it possible in particular to increase sedimentation and therefore, after possible filtration, to improve the pretreatment of the load.

Said optional pretreatment step a0) thus makes it possible to obtain a pretreated feed which then feeds the hydrogenation step a).

(a) Hydrogenation step

According to the invention, the process comprises a hydrogenation step a) 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 feedstock, optionally pretreated, optionally mixed with at least part of the second liquid effluent from step d) and a first gas stream comprising hydrogen, said hydrogenation reaction section being carried out at an average temperature between 140 and 400°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. And an hourly volume velocity between 0.1 and 10.0 h ^-1 , to obtain a hydrogenated effluent.

Step a) is in particular carried out under conditions of hydrogen pressure and temperature making it possible to carry out the hydrogenation of diolefins and olefins at the start of the hydrogenation reaction section while allowing, by a rising profile of the temperature d 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 c). 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 c) of hydrotreatment. In addition, the conditions of step a) make it possible to convert at least part of the chlorine, and preferably all of the chlorine.

Those skilled in the art easily understand that, in step a) of hydrogenation, hydrogenation reactions such as described above are carried out but in parallel also part of the other hydrotreatment reactions, and in particular hydrodesulfurization and hydrodenitrogenation, even if these reactions are rather favored in hydrotreatment step c) being generally carried out at higher temperature.

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 second liquid effluent from step d) and/or at least part of one or more cuts resulting from step g), in particular by the recycling rate and/or by the temperature of the recycled effluent.

The temperature difference between the inlet and the outlet of the reaction section of step a) means injection of a gaseous (hydrogen) or liquid cooling flow, in particular preferably a part of the second liquid effluent from step d).

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.).

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 (VVH) 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:

with T _inlet : the temperature of the flow entering the reaction section and T _outlet : 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 (VVH) is defined here as the ratio between the hourly volumetric flow rate of the charge comprising the 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 resulting from step d) recycled, at 15°C (in normal m ³ , noted Nm ³ , of H ₂ per m ³ of charge).

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 ).

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 one 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 VIB, preferably chosen from molybdenum and tungsten. According to this variant, the total content expressed in oxides of the metallic elements of groups VIB and VIII is preferably between 1% and 40% by weight, preferably from 5% to 30% by weight relative to the total weight of the catalyst. When the metal is cobalt or nickel, the metal content is expressed as CoO and NiO respectively. When the metal is molybdenum or tungsten, the metal content is expressed as MoO ₃ and WO ₃ 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 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 MoO ₃ 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 P ₂ O ₅ 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 B ₂ O ₃ 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 γ (gamma) or η (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 MoO ₃ 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 VIB 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). .

According to a variant, the feed, before being preferably mixed with at least part of the second effluent from step d), can be preheated by direct heating to a temperature of up to 200°C, preferably up to '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 charge to a temperature above 150°C, preferably above 180°C, and particularly preferably above 200°C is preferably carried out by indirect heating with a hot diluent, preferably by at least part of the second effluent from step d). 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. This helps limit high temperatures locally. This type of heating by mixing with an inert hot liquid therefore makes it possible to limit undesirable reactions such as the polymerization of diolefins (gum formation) and/or the formation of coke, and to adjust the entry temperature of the flow into the step a) so as to initiate the hydrogenation reaction of the unsaturations, preferably at the lowest possible temperature, while controlling the exotherm of these reactions by a dilution effect of the reactive species.

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

The energy necessary for the reaction and more precisely, the adjustment of the minimum temperature necessary for the activation of the saturation reactions of the double bonds is therefore advantageously achieved by mixing, upstream of step a), said charge comprising a pyrolysis oil and a hydrogen-rich gas, with a diluent and preferably a part of the second liquid effluent from separation step d), having optionally undergone a temperature adjustment and preferably having been either preheated or cooled and particularly preferred manner having been preheated.

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 e). At least part of this hydrogen-rich gaseous effluent coming from the hydrogen makeup and/or the gaseous effluent from separation step e) is advantageously injected mixed with at least part of the second effluent from step d) or separately, upstream of step a).

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, and halogens, in particular in chlorine. 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 or even nitrogen. Preferably, at least 50%, and more preferably at least 75%, of the chlorine and silicon of the initial charge are respectively eliminated during step a). Generally, the silicon content is less than 10 ppm by weight. The effluent obtained at the end of step a) of hydrogenation is sent, preferably directly, to step b) of hot separation.

(b) First hot separation step

According to the invention, the treatment process comprises a separation step b), supplied by the hydrogenated effluent from step a), said step being carried out at a temperature higher than the precipitation temperature of the ammonium halides and at a pressure substantially identical to the pressure of step a) to obtain at least a first gaseous effluent and a first liquid effluent.

The high temperature of this separation step b) avoids the precipitation of salt of ammonium halides which form by reaction between the halide ions and the ammonium ions in order to recover a first gaseous effluent containing the majority of the halides in the form hydrogen halides (HCl) and gaseous ammonia, and a first liquid effluent containing very little halides and ammonia.

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

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

By “pressure substantially equal to the pressure of step a)” is meant the pressure of step a) 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 0.3 MPa relative to the pressure of step a). Preferably, the pressure of step b) is the pressure of step a) reduced by pressure losses.

Separation step b) 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 several stripping column(s), this or these separator(s) (balloon(s)) and/or columns optionally being able to be supplied with a stripping gas, for example a gas flow rich in hydrogen. Preferably, step b) is implemented with a single separator (balloon).

Separation step b) 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 b) preferably uses a so-called “high pressure hot” separator, the pressure being substantially equal to the operating pressure of step a).

Gas/liquid separation has an efficiency corresponding to solubility and Henry's law. This means that an equilibrium quantity of halides in the form of hydrogen halides will remain in the first liquid effluent obtained in step b). These halides will be released into the gaseous effluent during cold separation step e), then discharged by dissolution in the aqueous effluent.

The hot separation of step b) makes it possible to keep the ammonia and the halides in the gas phase and to obtain a first liquid effluent essentially free of halogenated compounds which can be sent to the hydrotreatment step c) in order to remove the remaining impurities without observing the deposition of halogenated compounds on the catalyst.

According to a first variant, the first liquid effluent is sent, in part or in whole, and preferably in full, to the hydrotreatment step c).

According to another variant, the first liquid effluent can be sent to an additional cold separation/washing step b') before sending it to the hydrotreatment step c).

In this case, the process according to the invention comprises a separation step b'), supplied by at least part of the first liquid effluent from step b) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of the ammonium halides and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent, an aqueous effluent and a hydrocarbon effluent. In this way, the halides in the form of hydrogen halides dissolved in the first liquid effluent are released into the gaseous effluent during cold separation step b') then dissolved in the aqueous effluent.

This separation step b') of cold separation/washing is essentially operated in the same way as that of separation step e) described below, the separation conditions (pressure and temperature) may or may not be identical.

(c) Hydrotreatment step

According to the invention, the treatment method comprises a hydrotreatment step c) implemented in a hydrotreatment 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 hydrotreatment catalyst, said hydrotreatment reaction section being supplied at least by said first liquid effluent from step b) and a second gas stream comprising hydrogen, said reaction section hydrotreatment being carried out at an average temperature between 250 and 430°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. and an hourly volume velocity between 0.1 and 10.0 h ^-1 , to obtain a hydrotreated effluent.

Advantageously, step c) 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. In addition, the hydrogenation of the remaining olefins and halogenated compounds as well as the hydrodemetallation can continue even if most and preferably all of these impurities have been removed during step a).

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 (VVH) 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 d) is advantageously between 100 and 1500 Nm ³ of hydrogen per m ³ of fresh feed which feeds step a), and preferably between 200 and 1000 Nm ³ of hydrogen per m ³ of fresh feed which feeds step a), preferably between 250 and 800 Nm ³ of hydrogen per m ³ of fresh feed which feeds step a). The definitions of mean temperature (WABT), VVH and hydrogen coverage correspond to those described above.

Said hydrotreatment reaction section is supplied at least by said first liquid effluent from step b) and a second gas stream comprising hydrogen, advantageously at the level of the first catalytic bed of the first reactor in operation.

Advantageously, said step c) 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 c) 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 c) which comprises a single fixed bed reactor.

Advantageously, said hydrotreatment catalyst used in said step c) 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 VIB, preferably chosen from the group made up 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 MoO ₃ and WO ₃ 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 c) 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 MoO ₃ 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 optionally boron. When the phosphoric anhydride P ₂ O ₅ 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 B ₂ O ₃ 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 γ (gamma) or η (eta) alumina.

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

Advantageously, said hydrotreatment catalyst of the process has a specific surface area greater than or equal to 250 m ² /g, preferably greater than or equal to 300 m ² /g. The specific surface area of 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 furanic cycle or even sugars.

Preferably, step c) 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 c) allows the hydrogenation of at least 80%, and preferably all of the possible olefins remaining after hydrogenation step a) and any halogenated compounds remaining in the first effluent. liquid resulting from separation step b), but also the conversion at least in part of other impurities present in the feed, such as aromatic compounds, metal compounds, sulfur compounds, nitrogen compounds, oxygenated compounds. Preferably, the nitrogen content at the outlet of step c) is less than 100 ppm by weight, and preferably less than 10 ppm by weight. Preferably, the sulfur content at the outlet of step c) is less than 100 ppm by weight, and preferably less than 10 ppm by weight. Step c) 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 c) is less than 10 ppm by weight, and preferably less than 2 ppm by weight, and the silicon content is less than 5 ppm by weight. Preferably, the halogen element content at the output of step c) is less than 5 ppm by weight.

Depending on the content of sulfur compounds in the initial charge to be treated, a stream containing a sulfurizing agent can be injected upstream of step a) of hydrogenation and/or step c) of hydrotreatment and/or in upstream of one of the hydrocracking steps when they are present, preferably upstream of step a) of hydrogenation and/or step c) of hydrotreatment in order to ensure a sufficient quantity of sulfur to form or maintain the active species of the catalyst (in sulfur form). This activation or 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 H ₂ S gas, elemental sulfur, CS ₂ , 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 charges with a view to sulphurizing 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.

(d) Second hot separation step

According to the invention, the treatment process comprises a separation step d), supplied by the hydrotreated effluent from step c), said step being carried out at a temperature higher than the precipitation temperature of the ammonium halides formed. and at a pressure substantially identical to the pressure of step c) to obtain at least a second gaseous effluent and a second liquid effluent, part of which is optionally recycled upstream of step a).

This second hot separation step d) is essentially operated in the same way as that of the first separation step b), the separation conditions (pressure and temperature) being able to be identical or not.

Although the vast majority of halogenated compounds were eliminated during the first hot separation step b), the high temperature of this separation step d) avoids the precipitation of ammonium halide salts which form by reaction between the halide ions and ammonium ions formed during the hydrotreatment step of the remaining halogenated compounds. The high temperature also makes it possible to eliminate the NH ₃ and H ₂ S formed during the hydrotreatment step of the nitrogen and sulfur compounds respectively by said second gaseous effluent.

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

By “pressure substantially equal to the pressure of step c)” is meant the pressure of step c) 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 0.3 MPa relative to the pressure of step c). Preferably, the pressure of step d) is the pressure of step c) reduced by pressure losses.

Separation step d) 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 several stripping column(s), this or these separator(s) (balloon(s)) and/or columns optionally being able to be supplied with a stripping gas, for example a gas flow rich in hydrogen. Preferably, step d) is implemented with a single separator (balloon).

Separation step d) 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 d) preferably uses a so-called “high pressure hot” separator, the pressure being substantially equal to the operating pressure of step c).

Gas/liquid separation has an efficiency corresponding to solubility and Henry's law. This means that an equilibrium quantity of halides in the form of hydrogen halides will remain in the second liquid effluent obtained in step d). These halides will be released into the gaseous effluent during cold separation step e), then discharged by dissolution in the aqueous effluent.

Advantageously, part of the second liquid effluent from step d) is recycled upstream of step a).

Mixing the feedstock and the second liquid effluent (hot recycle) upstream of step a) makes it possible, on the one hand, to dilute the impurities in the feedstock and, on the other hand, to control the temperature in step a) in which highly exothermic reactions (notably the hydrogenation reactions of olefins and diolefins) occur. Mixing the feedstock and part of the second liquid effluent (hot recycle) upstream of step a) also allows indirect heating of the feedstock, by simple mixing of the “cold” feedstock with the hot recycle).

The quantity of the second liquid effluent from step d) recycled is adjusted so that the weight ratio between the recycle stream from step d) and the load comprising a pyrolysis oil, that is to say the load 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 the second liquid effluent from step d) 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 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.

According to a preferred variant, at least part of the second liquid effluent from step d) can advantageously be either cooled, or pre-heated, if necessary, or kept at the same temperature as at the outlet of step d) separation, before being preferably recycled upstream of step a) of hydrogenation, depending on the temperature and the flow rate of feed and hydrogen, so that the temperature of the incoming flow, comprising said feedstock preferably in mixture with at least part of said second liquid effluent from step d) and a gas rich in hydrogen, i.e. between 140 and 400°C, preferably between 220 and 350°C, and particularly preferred between 260 and 330°C.

High pressure and high temperature separation makes it possible, among other things, to maximize energy recovery by hot recycling of part of the second liquid effluent. Indeed, the energy to reach the entry temperature necessary in step a) is at least partly provided by the heat of part of the second liquid effluent from step d) and also makes it possible to reduce or even eliminate possible preheating by directly heating the load above a temperature above 200°C to avoid the formation of gums. In addition, the fact of preferably recycling at least part of the second liquid effluent at high pressure saves energy for its pressurization in step a).

(e) Cold separation step and washing

According to the invention, the treatment method comprises a separation step e), supplied by the first and second gaseous effluent and at least part of the second liquid effluent from step d) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of the ammonium halides and at a pressure substantially identical to or lower than the pressure of step d), to obtain at least a third gaseous effluent, an aqueous effluent and a hydrocarbon effluent.

The separation of step e) at a temperature lower than the precipitation temperature of the ammonium halides of the mixture of gaseous effluents and at least part of the second liquid effluent from step d) causes these halide salts to precipitate. 'ammonium. In addition, the hydrogen halides dissolved in the second liquid effluent from step d) are released and also form ammonium halides with the ammonia present in the gaseous effluent. Washing with an aqueous solution of this step e) makes it possible to dissolve these salts in the aqueous effluent. We thus obtain a hydrocarbon effluent freed from halides, a gaseous effluent freed from halides and an aqueous effluent in which the ammonium halide salts are dissolved.

The temperature at which the separation of step e) is carried out must be lower than the precipitation temperature of the ammonium halides in order to precipitate the ammonium halide salts. The temperature at which the separation of step e) 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.

By “pressure substantially equal to the pressure of step d)” is meant the pressure of step d) 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 0.3 MPa relative to the pressure of step d). Preferably, the pressure of step e) is the pressure of step d) reduced by pressure losses. The fact of carrying out at least part of the separation step e) at a pressure substantially identical to the operating pressure of step d) also facilitates the recycling of hydrogen contained in the gaseous effluent.

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

The separation step e) may also comprise a (first) separation step at a pressure substantially equal to the operating pressure of step d), followed by at least one other separation step carried out at an identical or lower temperature. and at a lower pressure in each separation step from the previous step e).

Separation step e) 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 several stripping column(s), this or these separator(s) (balloon(s)) and/or columns optionally being able to be supplied with a stripping gas, for example a gas flow rich in hydrogen. Preferably, step c) is implemented with a single separator (balloon).

The separation step e) is preferably implemented in at least one so-called high pressure or medium pressure separator flask at low temperature, also known to those skilled in the art under the name CHPS (for “Cold High Pressure Separator” according to Anglo-Saxon terminology). Thus, this step e) preferably uses a so-called “high pressure cold” separator, the pressure being substantially equal to the operating pressure of step d).

The washing/separation section of step e) 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 e) may for example include a column for stripping acidic water (also called a "sour water stripper" according to Anglo-Saxon terminology) from the withdrawn aqueous effluent, a column for washing acidic gases for purify the hydrogen-rich gas before recycling, a column for stabilizing the washed liquid effluent to eliminate dissolved gases.

The gaseous effluent obtained at the end of step e) 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 towards the stages a) of hydrogenation and/or c) of hydrotreatment and/or towards one or more stages g) of hydrocracking when they are present, the system of recycling which may include a purification section.

The gaseous effluent may also be subject to additional separation(s) in order to recover 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 aqueous effluent obtained at the end of step e) advantageously comprises dissolved ammonium salts and/or hydrogen halides (for example hydrochloric acid).

Depending on the content of halogenated compounds in the initial feed to be treated, a flow containing an amine such as for example monoethanolamine, diethanolamine and/or monodiethanolamine can be injected upstream of step a) of hydrogenation and/or or from step c) of hydrotreatment in order to ensure a sufficient quantity of ammonium ions to combine the halide ions formed during the hydrogenation and/or hydrotreatment steps, thus making it possible to limit the formation of hydrochloric acid and thus limiting corrosion downstream of the separation section. This injection may in particular be necessary when the load contains a lot of halogenated compounds or few nitrogenous compounds.

In a possible embodiment of the invention, separation step e) comprises the injection of an aqueous solution into the mixture of gaseous effluents and at least part of the second liquid effluent resulting from step d. ), followed by the washing/separation section advantageously comprising a separation phase making it possible to obtain at least one aqueous effluent loaded with dissolved ammonium halide salts, a washed liquid hydrocarbon effluent and a partially washed gaseous effluent. The aqueous effluent loaded with dissolved ammonium halide salts and the washed liquid hydrocarbon effluent can then be separated in a settling flask in order to obtain said hydrocarbon effluent and said 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 hydrocarbon effluent, which makes it possible to eliminate at least in part, preferably all of the hydrogen halides (of the HCl type) contained in the partially washed gaseous effluent and thus obtaining said gaseous effluent, preferably comprising essentially hydrogen, and an acidic aqueous stream. Said aqueous effluent from the settling tank can optionally be mixed with said acidic aqueous flow, and be used, optionally mixed with said acidic aqueous flow in a water recycling circuit to supply step e) of separation with said aqueous solution 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.

The hydrocarbon effluent from separation step e) can be sent, in part or in whole, either to a fractionation step f), or directly to a steam cracking step h), or even directly to a steam cracking unit. fuel storage.

According to another embodiment, part of the hydrocarbon effluent from separation step e) is recycled upstream of step a) and/or upstream of step c).

Mixing the feedstock and the hydrocarbon effluent from step e) (hot recycle) upstream of step a) and/or step c) makes it possible, on the one hand, to dilute the impurities in the load and on the other hand to control the temperature in steps a) and/or c) in which highly exothermic reactions occur. Mixing the feedstock and part of the hydrocarbon effluent from separation step e) (hot recycle) upstream of step a) also allows indirect heating of the feedstock, by simple mixing of the feedstock " cold” with hot recycle).

The recycling of at least part of the hydrocarbon effluent from step e) of separation also has the advantage that the first gaseous effluent from step b) (naphtha cutting) which has short-circuited the step c) of hydrotreatment and which is contained in the hydrocarbon effluent resulting from step e) of separation can be subjected by this recycling to hydrotreatment, in particular when at least part of the hydrocarbon effluent resulting from step e) separation is introduced upstream of step c) hydrotreatment.

The quantity of the hydrocarbon effluent from step e) of recycled separation is adjusted so that the weight ratio between the recycle stream from step e) and the feed comprising a pyrolysis oil, i.e. -say the load 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 the hydrocarbon effluent from the recycled separation step e) 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.

When part of the second liquid effluent from step d) is recycled upstream of step a) and part of the hydrocarbon effluent from separation step e) is recycled upstream of step a) and/or upstream of step c), the weight ratio between the total recycle stream (effluent from step d) and effluent from step e)) and the feed 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.

(f) Splitting step (optional)

The process according to the invention may comprise a step of fractionating all or part, preferably all, of the hydrocarbon effluent resulting from step e), to obtain

at least a fourth gaseous effluent, a naphtha cut and at least one middle distillate cut.

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

The term “middle distillate cut” means a hydrocarbon cut comprising compounds having a boiling point generally greater than 175°C. The heavy cut may include middle distillates such as a diesel cut and/or a kerosene cut. It can also include heavier compounds.

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.

Step f) makes it possible in particular to eliminate gases dissolved in the hydrotreated 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 1.0 MPa abs., preferably between 0.1 and 1.0 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 supplied by the liquid hydrocarbon effluent from step e) and by a flow of water vapor. The liquid hydrocarbon effluent from step e) can optionally 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 naphtha cut is advantageously withdrawn from the reflux flask. The middle distillate cut is advantageously withdrawn from the bottom of the stripping column.

According to another embodiment, step e) can be carried out in a section comprising one or more separator tank(s). We will advantageously use at least one so-called low pressure separator balloon at low temperature in order to eliminate the gases dissolved in the hydrocarbon liquid effluent.

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

The naphtha cut and the middle distillate cut, possibly 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 naphtha cut, all or part, is sent to a steam cracking unit, while the middle distillate cut 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 (generally comprising compounds having a boiling point less than or equal to 175°C, preferably between 80 and 175°C), and a middle distillate cut (generally comprising compounds having a boiling point greater than 175°C and less than 385°C), and a heavy hydrocarbon cut (generally comprising compounds having a boiling point boiling point greater than or equal to 385°C). The naphtha cut can be sent, in whole or in part, to a steam cracking unit and/or to the naphtha storage unit from conventional 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 feeds, or to step g) of hydrocracking when it is present, or still be recycled; the heavy cut can for its part be sent, at least in part, to a steam cracking unit, or be sent to the hydrocracking stage when 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 (generally comprising compounds having a boiling point less than or equal to 175°C, of preferably between 80 and 175°C), and a kerosene cut (generally comprising compounds having a boiling point greater than 175°C and less than or equal to 280°C), a diesel cut (generally comprising compounds having a boiling point boiling point greater than 280°C and less than 385°C) and a heavy hydrocarbon cut (generally comprising compounds having a boiling point greater than or equal to 385°C). The naphtha cut, 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. Diesel and/or kerosene cuts can also be sent to hydrocracking step g) when present. 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 naphtha cut resulting from step f) is fractionated into a heavy naphtha cut (generally comprising compounds having a boiling point between 80 and 175°C) and a light naphtha cut (generally comprising compounds having a boiling point below 80°C), at least part of said heavy naphtha cut being sent to an aromatic complex comprising at least one naphtha reforming step in order 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.

When at least one of these cuts is recycled, it is advantageously recycled upstream of step a) and/or upstream of step c).

The quantity of the recycled cut(s) is adjusted so that the weight ratio between the recycle flow and the feed comprising a 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 the recycled cut(s) is adjusted so that the weight ratio between the recycle flow and the charge comprising a plastic pyrolysis oil is between 0.01 and 10, preferably between between 0.1 and 7, and particularly preferably between 0.2 and 5.

When one or more cuts from step f) of fractionation are recycled and/or part of the second liquid effluent from step d) and/or part of the hydrocarbon effluent from the step e) of separation, the weight ratio between the total recycle flow and the charge comprising a plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 and 7, and particularly preferably included between 0.2 and 5.

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.

(g) Hydrocracking stage (optional)

According to a variant, the process of the invention may comprise a step g) of hydrocracking carried out after step c) of hydrotreatment with at least part of said hydrotreated effluent from step c) or carried out after step f) fractionation with at least a portion of middle distillate cutting.

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 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), we can transform these compounds at least in part into light compounds by hydrocracking, a cut generally favored for a unit of steam cracking.

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 hydrotreated effluent from step c) and/or by at least part of the middle distillate cut from step f) and a third gas stream comprising hydrogen, said hydrocracking reaction section being carried out at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly 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 480°C, preferably between 320 and 450°C, at a partial pressure of hydrogen between 1.5 and 20.0 MPa abs ., preferably between 3 and 18.0 MPa abs, and at an hourly volume velocity (VVH) 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), VVH 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 c) 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 hot separation and/or in step c) of hydrotreatment and/or in the step d) of hot separation and/or in step e) of cold separation and/or in step f) of fractionation.

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') 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 a 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, and 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 per pass in the second stage is kept moderate in order to maximize the selectivity to compounds of the naphtha cut (having a boiling point less than or equal to 175°C, in particularly between 80 and less than or equal to 175°C). 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), 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 step a) of hydrogenation and/or in step b) of hot separation and/or in step c) hydrotreatment and/or in step d) of hot separation and/or in step e) of cold separation and/or in step f) of fractionation.

The hydrocracking step(s) therefore does not necessarily make it possible to transform all the hydrocarbon compounds in the middle distillate cut into hydrocarbon compounds having a boiling point less than 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. 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 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 VIB of the periodic table and/or at least one metal from group VIII.

Preferably, the hydrocracking catalyst(s) comprise a hydro-dehydrogenating 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 metal from Group VIB 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 VIB 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 MoO ₃ and WO ₃ 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 VIIA (chlorine, preferred fluorine), optionally at least one element from group VIIB (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 VIB metal and optionally at least one non-noble Group VIII metal, at least one promoter element, and preferably phosphorus, 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 furanic 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 VIB 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 implying a reduction step in its preparation.

The gas stream comprising hydrogen, which feeds the hydrogenation, hydrotreatment and hydrocracking reaction section may consist of additional hydrogen and/or recycled hydrogen resulting in particular from step e) separation or even from step f) of fractionation. 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 hydrotreated effluent or said hydrocarbon cut(s) thus obtained by treatment according to the process of the invention of a pyrolysis oil, has a composition compatible with the specifications of an input charge to a steam cracking unit. In particular, the composition of the hydrotreated effluent or said hydrocarbon cut(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, 8 ppm 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
an iron element content (Fe) less than or equal to 200 ppb by weight,
- the sulfur content is less than or equal to 500 ppm by weight, preferably less than or equal to 200 ppm by weight,
- the nitrogen content is less than or equal to 100 ppm by weight, preferably less than or equal to 50 ppm by weight and preferably less than or equal to 5 ppm by weight
- the total chlorine element content is less than or equal to 5.0 ppm by weight, preferably less than 1.0 ppm by weight,
- the content of olefinic compounds (mono- and di-olefins) is less than or equal to 5.0% by weight, preferably less than or equal to 2.0% by weight, preferably less than or equal to 0.1% by weight,
- the mercury content is less than or equal to 5 ppb by weight, preferably less than 3 ppb by weight.

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 the pyrolysis oils to obtain at least one effluent which can be injected, in whole or in part, into at least one steam cracking unit.

Heavy metal adsorption step (optional)

Any gaseous effluent and/or any liquid effluent resulting from at least one of steps b), b'), d) and e) of separation or from step f) of fractionation, can be subjected to an optional step of adsorption of heavy metals.

The gaseous effluents are in particular the first gaseous effluent from step b) and/or the gaseous effluent from step b') and/or the second gaseous effluent from step d) and/or third effluent gas from step e) and/or the fourth gaseous effluent from step f).

The liquid effluents are in particular the first liquid effluent from step b) and/or the hydrocarbon effluent from step b') and/or the second liquid effluent from step d), the hydrocarbon effluent from step e) and/or at least one of the cuts 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 may be present in the feed and/or formed during the process steps. Their elimination or reduction may in particular be necessary when at least part of said gaseous and liquid effluents is intended to be sent to a steam cracking stage, either directly or after having undergone one or more optional additional stages such as the fractionation stage. f). The specifications for metal impurities entering the steam cracking stage, particularly mercury, may require such a step.

This optional adsorption step 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 ppm by weight of metallic elements of heavy metals (As, Zn, Pb, Hg , …), and in particular when at least one of these effluents or the load respectively comprises more than 10 ppm by weight of mercury, more particularly more than 15 ppm by weight of mercury.

Said optional adsorption step is advantageously carried out at a temperature between 20 and 150°C, preferably between 40 and 100°C, and at a pressure between 0.15 and 10.0 MPa abs, preferably between 0. 2 and 1.0 MPa abs.

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 based on sulfur in the elemental form or in the form of metal sulfide.

The porous support can be chosen indifferently from the families of aluminas, silica-aluminas, silicas, zeolites, 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 the elemental form or in the form of metal sulfide, in particular a sulfide 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%. 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 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, a mode of operation 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. Another mode of operation is to have at least two columns operating in series in swappable mode.

(h) Steam cracking step (optional)

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

Advantageously, the gaseous effluent(s) resulting from step b), b'), d), e) of separation and/or f) of fractional and containing ethane, propane and butane, can (can) in whole or in 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 second (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 (or fractionation). 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 middle distillate cut.

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 with a view to recovering at least a light fraction comprising hydrogen and compounds having 2 to 5 carbon atoms, and a fraction comprising pyrolysis gasoline, and possibly a larger fraction. heavy. 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.

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:

Description Methods Density @15°C ASTM D4052 Sulfur Content ISO 20846 Nitrogen Content ASTM D4629 Acid number ASTM D664 Bromine Index ASTM D1159 Maleic anhydride index MAV method (1) Oxygen content Combustion + Infra-red Paraffin content UOP990-11 Content of Naphthenes and Olefins UOP990-11 Aromatic content UOP990-11 Halogen content ASTM D7359 Chlorine content ASTM D7536 Metal Content: ASTM D5185 P Fe If N / A B Simulated distillation ASTM D2887

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

List of Figures

The mention of the elements referenced in Figures 1 to 4 allows a better understanding of the invention, without it being limited to the particular embodiments illustrated in Figures 1 to 4. The different embodiments presented can be used alone or in combination with each other, without limitation of combination. In the figures, the same references designate identical or similar elements.

[Fig 1]

There represents the diagram of a particular embodiment of the process of the present invention, comprising:

- a step a) of hydrogenation of a hydrocarbon feedstock resulting from the pyrolysis of plastics 1 preferably mixed with at least part of the second recycled liquid effluent 11a resulting from step d) and in the presence of a rich gas in hydrogen 2 and optionally an amine supplied by stream 3 and optionally a sulfur compound supplied by stream 4, carried out in at least one fixed bed reactor comprising at least one hydrogenation catalyst, to obtain a hydrogenated effluent 5;

- a step b) of separation of the hydrogenated effluent 5 carried out at high pressure and high temperature (HHPS) to obtain at least a first gaseous effluent 6, and a first liquid effluent 7,;

- a step c) of hydrotreating the first liquid effluent 7 from step b), in the presence of hydrogen 8 carried out in at least one fixed bed reactor comprising at least one hydrotreatment catalyst, to obtain a hydrotreated effluent 9;

- a step d) of separation of the hydrotreated effluent 9 carried out at high pressure and high temperature (HHPS) to obtain at least a second gaseous effluent 10, and a second liquid effluent 11 of which at least part 11a is preferably recycled in upstream of step a);

- a separation step e) carried out at high pressure and low temperature (CHPS) and supplied by the first gaseous effluent 6, the second gaseous effluent 10 and at least part of the second liquid effluent 11 from step d) and a aqueous solution 12 and making it possible to obtain at least a third gaseous effluent 13 comprising hydrogen, an aqueous effluent 14 containing dissolved salts, and a hydrocarbon effluent 15;

- optionally a step f) of fractionating the hydrocarbon effluent 15 making it possible to obtain at least a fourth gaseous effluent 16, a naphtha cut 17 (generally comprising compounds having a boiling point less than or equal to 175°C) and a middle distillate cut 18 (generally comprising compounds having a boiling point above 175°C).

At the end of step f), part of the naphtha cut 17 can be sent to a steam cracking process (not shown). Another part of the naphtha cut 17 can feed the hydrogenation step a) and/or the hydrotreatment step c) (not shown).

[Fig 2]

There represents the diagram of another particular embodiment of the method of the present invention which is based on the diagram of the . This diagram includes in particular a hydrocracking step g) after the fractionation step f) in which at least part of the middle distillate cut 18 from step f) feeds this hydrocracking step g) which is carried out in at least one fixed bed reactor comprising at least one hydrocracking catalyst and is supplied with hydrogen 19. The hydrocracked effluent 20 is recycled upstream of separation step e).

[Fig 3]

There represents the diagram of another particular embodiment of the method of the present invention which is based on the diagram of the . This diagram includes in particular a hydrocracking step g) directly after the hydrotreatment step c) in which at least part of the hydrotreated effluent 9 from step c) feeds this hydrocracking step g) which is carried out in at least one fixed bed reactor comprising at least one hydrocracking catalyst and is supplied with hydrogen 19. Sending the hydrotreated effluent from hydrotreatment step c) directly into a hydrocracking step has the advantage of not having to repressurize it and therefore saves energy. The hydrocracked effluent 20 is then sent to hot separation step d).

[Fig 4]

There represents the diagram of another particular embodiment of the method of the present invention which is based on the diagram of the . This diagram includes in particular a cold separation/washing step b') after the hot separation step b) and before the hydrotreatment step c) in which at least part of the first liquid effluent 7 from the step b) feeds this step b'). Step b') is carried out at high pressure and low temperature (CHPS) and is supplied by the first liquid effluent 7 and an aqueous solution 21. It makes it possible to obtain at least one gaseous effluent 22 comprising hydrogen, a aqueous effluent 23 containing dissolved salts, and a hydrocarbon effluent 24 which is introduced into hydrotreatment step c). In this way, the halides in the form of hydrogen halides dissolved in the first liquid effluent 7 are released into the gaseous effluent 22 during the cold separation step b').

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 c) of hydrotreatment, at the inlet of step e) 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 to 4, 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 is an example not in accordance with the invention without the first step b) of hot separation between step a) of hydrogenation and step c) of hydrotreatment.

Example 2 is an example in accordance with the invention with the first step b) of hot separation between step a) of hydrogenation and step c) of hydrotreatment.

Charge 1 treated in the process is a plastic pyrolysis oil (i.e. comprising 100% by weight of said plastic pyrolysis oil) having the characteristics indicated in Table 2.

Table 2: load characteristics

Description Methods Unit Charge Density @15°C ASTM D4052 g/cm ³ 0.835 Sulfur Content ISO 20846 ppm weight 135 Nitrogen Content ASTM D4629 ppm weight 2109 Acid number ASTM D664 mg KOH/g 1.5 Bromine index ASTM D1159 g/100g 59 Maleic anhydride index MAV method mg/100g 36 Olefin content UOP990-11 % weight 25 Chlorine content ASTM D7536 ppm weight 500 Silicon content: ASTM D5185 ppm weight 92 Simulated distillation ASTM D2887 0% °C 40 10% °C 98 30% °C 161 50% °C 232 70% °C 309 90% °C 394 100% °C 432

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 operating conditions indicated in Table 4 allowing to obtain a hydrogenated effluent 5.

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

Table 3: species conversions during hydrogenation step a)

Chlorine conversion rate % >90 Conversion rate of diolefins % >70 Olefin conversion rate % >70 Silicon retention % > 85

According to example 1 (non-compliant), the hydrogenated effluent 5 from step a) of hydrogenation is subjected directly, without separation, to a step c) of hydrotreatment carried out in a fixed bed and in the presence of hydrogen 8, and a NiMo type hydrotreatment catalyst on alumina under the conditions presented in Table 4. The hydrotreated effluent is then subjected to a hot separation step d) at a pressure substantially identical to that of the step c) and whose temperature is controlled at 300°C, making it possible to obtain a gaseous effluent and a liquid effluent. The gaseous effluent and the liquid effluent are then subjected to a cold separation step e): a flow of water is injected into the mixture, the final mixture reaches a temperature of 40°C in a cold HP tank operating at a pressure substantially identical to that of step d), at the outlet of which we obtain a gas fraction rich in hydrogen, an aqueous fraction and the washed hydrocarbon effluent.

The hydrotreated effluent is then introduced into a so-called low pressure separator tank at low temperature in order to eliminate the dissolved gases. The properties of the hydrotreated effluent, free of dissolved gases, are presented in Table 4 and comply with the specifications of a feed for a steam cracker. All or part of the hydrotreated effluent obtained can then be upgraded in a steam cracking step to form olefins which can be polymerized to form recycled plastics.

According to Example 2 (compliant), the hydrogenated effluent 5 from step a) of hydrogenation is subjected to a step b) of separation (hot) at a pressure substantially identical to that of step a) and whose temperature is controlled at 300°C, making it possible to obtain a first gaseous effluent 6 and a first liquid effluent 7.

The first liquid effluent 7 is then introduced into hydrotreatment step c) carried out in a fixed bed and in the presence of hydrogen 8, and a NiMo type hydrotreatment catalyst on alumina under the conditions presented in Table 4 to obtain a hydrotreated effluent 9.

The hydrotreated effluent 9 is then introduced into the second hot separation step d) at a pressure substantially identical to that of step c) and whose temperature is controlled at 300°C, making it possible to obtain a second gaseous effluent 10 and a second liquid effluent 11.

The first gaseous effluent 6, the second gaseous effluent 10 and the second liquid effluent 11 are then subjected to a cold separation step e): a flow of water 12 is injected into the mixture, the final mixture reaches a temperature of 40 °C in a cold HP tank operating at a pressure substantially identical to that of step d), at the outlet of which we obtain a hydrogen-rich gas fraction 13, an aqueous fraction 14 and the washed hydrocarbon effluent 15.

The hydrotreated effluent 15 is then introduced into a so-called low pressure separator tank at low temperature in order to eliminate the dissolved gases. The properties of the hydrotreated effluent, free of dissolved gases, are presented in Table 4 and comply with the specifications of a feed for a steam cracker. All or part of the hydrotreated effluent obtained can then be upgraded in a steam cracking step to form olefins which can be polymerized to form recycled plastics.

Table 4: operating conditions of steps a), b), c), d) and e) and properties of the hydrotreated effluent.

Example 1
(improper) Example 2 (compliant) Hydrogenation step a) Average temperature (WABT) °C 280 280 ppH ₂ at the outlet of a) MPa 5.2 5.2 Total pressure at the outlet of a) MPa 6.1 6.1 H ₂ /HC ^Nm3 / ^m3 300 300 VVH h ^-1 0.64 0.64 Step b) hot separation T °C 300 Total pressure at the outlet of b) MPa 5.8 Step c) hydrotreatment Average temperature (WABT) °C 302 294 _ppH2 MPa 5.2 5.2 Total pressure at the outlet of c) MPa 5.9 5.9 H ₂ /HC ^Nm3 / ^m3 300 300 VVH h ^-1 0.76 0.76 Step d) hot separation T °C 300 300 Total pressure at the outlet of d) MPa 5.6 5.6 Step e) cold separation T °C 300 300 Total pressure at the outlet of e) MPa 5.0 5.0 Properties of hydrotreated effluent N content ppm wt <5 <5 S content mg/kg <10 <10 Cl content mg/kg <5 <5 Metal Content mg/kg < 1 < 1 Si content mg/kg < 1 < 1 Hg content ppb <5 <5

It is observed that to obtain a hydrotreated effluent meeting the specifications of a feed for a steam cracker, the average temperature of hydrotreatment step c) according to example 2 is 8°C lower than that according to example 1 .

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

When we set a temperature difference between the start of the cycle and the end of the catalytic cycle of 60°C, we can deduce a catalyst deactivation speed and the duration of the catalytic cycle.

Table 5 shows the effect on cycle time of hydrotreatment step c) with or without hot separation step b). We observe that the cycle duration according to example 1 is 12 months, while the cycle duration according to example 2 is 20 months.

Table 5: catalytic cycle duration of the hydrotreatment step with ∆T (SOR*-EOR**) = 60 °C

Example 1
(improper) Example 2 (compliant) Deactivation Speed °C/month 5 3 ∆T (SOR*-EOR**) °C 60 60 T _SOR °C 302 294 T _EOR °C 362 354 Cycle Duration month 12 20

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

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

The fact that the gaseous effluent from step b) according to example 2 short-circuits hydrotreatment step c) makes it possible to reduce the size of the reactor of hydrotreatment step c) and therefore saves money. in need of catalyst and equipment. This is demonstrated in Table 6.

Table 6: volume of catalyst necessary in hydrotreatment step c)

Example 1
(improper) Example 2 (compliant) Input flow c)/load inlet flow a) % 100 68.3 Volume of catalyst required m ³ 10.0 7.1

Claims (18)

Process for treating a load comprising a plastics and/or tire pyrolysis oil comprising halogenated compounds, said process comprising:
a) 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 catalyst hydrogenation, said hydrogenation reaction section being supplied at least by said feed possibly mixed with at least part of the second liquid effluent from separation step d) and a first gas stream comprising hydrogen, said hydrogenation reaction section being carried out at an average temperature between 140 and 400°C, a partial pressure of hydrogen between 1.0 and 10.0 MPa abs. And an hourly volume velocity between 0.1 and 10.0 h ^-1 , to obtain a hydrogenated effluent,
b) a separation step, supplied by the hydrogenated effluent from step a), said step being carried out at a temperature higher than the precipitation temperature of the ammonium halides and at a pressure substantially identical to the pressure of the step a) to obtain at least a first gaseous effluent and a first liquid effluent,
c) a hydrotreatment step 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 catalyst hydrotreatment, said hydrotreatment reaction section being supplied at least by said first liquid effluent from step b) and a second gas stream comprising hydrogen, said hydrotreatment reaction section being implemented at a temperature average 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,
d) a separation step, supplied by the hydrotreated effluent from step c), said step being carried out at a temperature higher than the precipitation temperature of the ammonium halides and at a pressure substantially identical to the pressure of the step c) to obtain at least a second gaseous effluent and a second liquid effluent, part of which is optionally recycled upstream of step a),
e) a separation step, supplied by the first and second gaseous effluent and at least part of the second liquid effluent from step d) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature ammonium halides and at a pressure substantially identical to or lower than the pressure of step d), to obtain at least a third gaseous effluent, an aqueous effluent and a hydrocarbon effluent,
f) optionally a step of fractionating all or part of the hydrocarbon effluent from step e), to obtain at least a fourth gaseous effluent, a naphtha cut and at least one middle distillate cut,
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 hydrotreated effluent from step c) and/or by at least part of the middle distillate cut from step f) and a third gas stream comprising hydrogen, said hydrocracking reaction section being carried out at an average temperature between 250 and 450°C, a partial pressure of hydrogen between 1.5 and 20.0 MPa abs. and an hourly volume velocity between 0.1 and 10.0 h ^-1 to obtain a first hydrocracked effluent. Process according to the preceding claim comprising the fractionation step f). Process according to one of the preceding claims comprising the hydrocracking step g). Method according to one of the preceding claims comprising a separation step b') carried out between steps b) and c), supplied with at least part of the first liquid effluent from step b) and an aqueous solution, said step being carried out at a temperature lower than the precipitation temperature of the ammonium halides and at a pressure substantially identical to or lower than the pressure of step b), to obtain at least one gaseous effluent, an aqueous effluent and a hydrocarbon effluent. Process according to the preceding claim, in which step b) or step d) is carried out at a temperature between 200 and 450°C and step e) or step b') is carried out at a higher temperature. or equal to 20°C and less than 200°C. Method according to one of the preceding claims, in which at least part of the second liquid effluent obtained in step d) is recycled upstream of step a). Method according to one of the preceding claims, in which at least part of the hydrocarbon effluent from step e) is recycled upstream of step a) and/or upstream of step c). Method according to one of the preceding claims, in which in step a) the hydrogen coverage is between 250 and 800 Nm ³ of hydrogen per m ³ of charge (Nm ³ /m ³ ). Method according to one of the preceding claims, comprising at least one step a0) of pretreatment of the pyrolysis oil fraction of plastics and/or tires, said pretreatment step being implemented upstream of step a) and comprises an adsorption step and/or a filtration step and/or a centrifugation step and/or a decantation step and/or an electrostatic separation step and/or a washing step using an aqueous solution and/or a gas stripping step. Method according to one of the preceding claims, in which the hydrocarbon effluent resulting from separation step e), or at least one of the cuts resulting from step f), is sent in whole or in part towards 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. Method according to one of the preceding claims in which the reaction section of step a) uses at least two reactors operating in switchable mode. Method according to one of the preceding claims in which a flow containing an amine and/or a sulfur compound is injected upstream of step a). Process 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 an element from group VIII and at least one element from group VIB, i.e. at least one element from group VIII. Process 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 hydro function. -dehydrogenating agent comprising at least one element from group VIII and/or at least one element from group VIB. Process according to one of the preceding claims, which 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 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. Process 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 hydro- dehydrogenating agent comprising at least one metal from Group VIB 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. Product obtained by the process according to one of claims 1 to 16. Product according to claim 17, which comprises, relative 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,
- a silicon element content less than or equal to 5.0 ppm by weight,
- a sulfur content less than or equal to 500 ppm by weight,
- a nitrogen content less than or equal to 100 ppm by weight,
- a chlorine element content less than or equal to 10 ppm by weight,
- a mercury content less than or equal to 5 ppb by weight.