TW202407084A - Process for the treatment of plastics pyrolysis oil including an h2s recycling stage - Google Patents

Abstract

The present invention relates to a process for the treatment of a plastics pyrolysis oil, comprising: - a hydrotreating of the feedstock in the presence of hydrogen and of a catalyst, - a separation/scrubbing of the hydrotreated effluent in the presence of an aqueous solution in order to obtain at least a first aqueous effluent and a hydrotreated hydrocarbon effluent, - a separation of the H2S contained in the first aqueous effluent, in order to obtain a gas phase containing the H2S and a second aqueous effluent, it being possible for said gas phase containing the H2S to be, at least partly, recycled upstream of stage (b), - a separation of the NH3 contained in the second aqueous effluent in order to obtain a gas phase containing NH3 and a third aqueous effluent. The present invention makes it possible, by the recycling of the H2S resulting from the process, to decrease the consumption of sulfiding agent for keeping the catalysts in the sulfide form in feedstocks containing only a little sulfur.

Description

Methods for treating plastic pyrolysis oil in the recovery stage containing H2S

The present invention relates to a method for treating plastic pyrolysis oil in order to obtain a hydrocarbon effluent which can be upgraded in units for storing petroleum, injection or gas oil fuels or used as feedstock for stream cracking units. More specifically, the present invention relates to a method for processing feedstock produced by the pyrolysis of plastic waste, which allows the H ₂ S-containing gas phase produced by the method to be recovered for use in the catalysis of the method. The catalyst in the form of sulfide is maintained during this stage and thus the consumption of sulfidizing agent to be added is reduced.

The plastic produced by the collection and storage channels can undergo pyrolysis stages in order to obtain inter alia pyrolysis oil. These plastic pyrolysis oils are typically incinerated to generate electricity and/or used as fuel in industrial or municipal heating boilers.

Another approach to upgrading plastic pyrolysis oils is to use these plastic pyrolysis oils as feedstock for steam cracking units to (re)produce olefins, which are the constituent monomers of certain polymers. However, plastic waste is usually a mixture of several polymers, such as polyethylene, polypropylene, polyethylene terephthalate, polyvinyl chloride or polystyrene. Furthermore, depending on the application, these plastics may contain other compounds in addition to the polymer, such as plasticizers, pigments, dyes or residues of polymerization catalysts. Plastic waste may additionally contain small amounts of biomass derived from, for example, household waste. The handling of waste can also cause corrosion on the one hand (especially storage, mechanical treatment, sorting or pyrolysis) and on the other hand (storage and transport of pyrolysis oil). As a result, the oil produced by the pyrolysis of plastic waste contains a large number of impurities, especially dienes, metals, especially iron, silicon, or halogen compounds, especially chlorine-based compounds, and miscellaneous elements such as sulfur and oxygen. and nitrogen, and insoluble materials, which are often present in high amounts and are incompatible with the steam cracking unit or units downstream of the steam cracking unit (especially polymerization processes and selective hydrogenation processes). Such impurities can create operability problems and in particular corrosion, coking or catalytic deactivation problems, or incompatibility problems with respect to the application of the target polymer. The presence of dienes can also lead to instability problems in pyrolysis oils, which are characterized by the formation of gums. The glue and insoluble materials that may be present in the pyrolysis oil can cause clogging problems in these processes.

Furthermore, during the steam cracking stage, the yield of light olefins (especially ethylene and propylene) sought for use in the petrochemical industry depends largely on the quality of the feedstock sent to the steam cracking. BMCI (Bureau of Mines Correlation Index) is often used to characterize hydrocarbon fractions. , this index developed for hydrocarbon products produced from crude oil is calculated by measuring density and average boiling point: it is equal to 0 for linear paraffins and 100 for benzene. Therefore, its value increases proportionally with the condensed aromatic structure of the product analyzed, with naphthenes having intermediate BMCI values between paraffins and aromatics. In general, the yield of light olefins increases when the paraffin content increases and therefore when the BMCI decreases. Conversely, as the BMCI increases, the yield of undesirable heavy compounds and/or coke increases.

Document WO 2018/055555 provides an overall method for recycling plastic waste, which is very versatile and relatively complex, ranging from each stage of pyrolysis of plastic waste up to the steam cracking stage. The method of application WO 2018/055555 includes in particular a stage of hydrotreating the liquid phase directly produced by pyrolysis, preferably under fairly stringent conditions, especially with respect to temperature, for example at a temperature of 260 to 300°C; A stage of separation of the hydrotreating effluent followed by hydrodealkylation of the separated heavy effluent at a preferably elevated temperature (eg 260 to 400°C).

Unpublished patent application FR 21/00.026 describes a method for treating plastic pyrolysis oil with the aim of reducing and/or removing the impurities contained in the pyrolysis oil in order to obtain an effluent compatible with steam crackers . The method includes the following stages: a) Hydrogenate the raw material in the presence of at least hydrogen and at least one hydrogenation catalyst at an average temperature of 140 to 340°C, the outlet temperature of stage a) being at least 15°C greater than the inlet temperature of stage a), in order to obtain a hydrogenation outflow thing; b) hydrotreating the hydrogenated effluent in the presence of at least hydrogen and at least one hydrotreating catalyst so as to obtain a hydrotreated effluent, the average temperature of stage b) being greater than the average temperature of stage a); c) Separating the hydrotreated effluent in the presence of an aqueous stream at a temperature of 50 to 370°C in order to obtain at least a gas effluent, an aqueous liquid effluent and a hydrocarbon liquid effluent.

Therefore, one approach for removing impurities contained in ex-plastic pyrolysis oil is to perform hydrotreating in the presence of a catalyst that is active in the sulfide form. In the context of feedstocks containing plastic pyrolysis oils, the feedstocks available are generally quite low in sulfur content. In fact, a minimum pH ₂ Sp is required in the hydrotreating reactor in order to keep the catalyst in the sulfide form and therefore not reduce it. In order to maintain a sufficient pH ₂ Sp in the reactor and in view of the fact that the feed does not contain sufficient sulfur, it is generally practical or even necessary to continuously add a sulfurizing agent, usually DMDS (dimethyl disulfide), to the feed. The sulfiding agent decomposes very rapidly by the reactor inlet temperature and the action of hydrogen to obtain H ₂ S and thus provides the amount of H ₂ S required to ensure a minimum and sufficient pH ₂ Sp.

After hydrotreating, at least a portion of the H ₂ S contained in the effluent forms ammonium sulfide salt ((NH ₄ ) ₂ S) with NH ₃ generated by hydrogenation of nitrogen compounds during hydrotreating. Unlike conventional feedstocks of the fossil type, ex-plastic pyrolysis oils generally contain a greater nitrogen content than sulfur content. These salts are generally removed by washing with water, followed by a (single) stage of steam stripping of the aqueous effluent, which allows to obtain a purified aqueous effluent and a gas phase containing H ₂ S and NH ₃ , which It is generally discharged together at the top of the stripper. Subsequently, the gas phase containing _H2S and _NH3 is typically incinerated to form _SOx (sulfur oxides) and _N2 or _NOx (nitrogen oxides).

The gas phase containing H ₂ S and NH ₃ can be recovered and sent to the inlet of the hydrotreating unit in order to maintain pH ₂ Sp in the reactor without adding sulfurizing agent. However, the _NH3 contained in the gas phase prevents this, since there is a concentration of _NH3 in the recovery loop that would be detrimental to the operation of the unit. Furthermore, the presence of NH ₃ lowers the pH ₂ p. Therefore, the gas phase containing H ₂ S and NH ₃ cannot be directly reused as an H ₂ S source to keep the catalyst in the sulfide form.

The present invention provides a method for treating feedstocks containing plastic pyrolysis oil, which allows the recovery of a phase containing only H ₂ S in order to use this phase by separation in two stages (generally by stripping). The H ₂ S source at the inlet of the catalytic unit enables separation of H ₂ S and NH ₃ . This is due to the use of two stripping columns operating under different operating conditions, which makes it possible to separate H ₂ S from NH ₃ and therefore recover: - a gas phase containing H ₂ S, which can be recovered at the inlet of the hydrotreating unit ; - and the gas phase containing _NH3 , which can be incinerated or also recovered at the inlet of the hydrotreating unit.

This two-stage stripping, which makes it possible to separate H ₂ S from NH ₃ , therefore exhibits the following advantages: - removal of NH ₃ from the phase containing only H ₂ S allows recovery of H ₂ S at the inlet of the catalytic unit of the process; - High degree of minimization of the consumption of vulcanizing agents; - Elimination of the NH ₃ -containing gas phase by incineration, which is easier since it no longer contains H ₂ S, which forms pollutants of the SO _x type; - NH ₃ -containing gas phase Recovery can also be carried out at the inlet of the catalytic unit, advantageously in a stoichiometric amount suitable for the formation of salts during separation/washing stage c); - better observation of the environmental limits regarding SO _x , since most of the H ₂ S is not After incineration (but on the contrary, recovered in the circuit); - Hydrogen consumption in the hydrotreating unit decreases since the sulfurizing agent (DMDS) consumes hydrogen for decomposition; - From the separation/washing section described below ( The top of stage c) completely removes _NH3 from the gas effluent containing hydrogen and/or light hydrocarbons. This is because the NH ₃ has been captured by the excess H ₂ S that has been recovered in the form of ammonium sulfide contained in the aqueous effluent. Therefore, the gas effluent thus released from _NH3 can be sent to a steam cracker in order to increase the overall yield of olefins.

More specifically, the invention relates to a method for treating feedstock containing plastic pyrolysis oil, which comprises: a) a hydrogenation stage optionally carried out in a hydrogenation reaction zone, using at least one catalytic bed with n A fixed bed reactor, n is an integer greater than or equal to 1, each catalytic bed contains at least one hydrogenation catalyst, and the hydrogenation reaction section is fed with at least the raw material and a gas stream containing hydrogen. The hydrogenation reaction section is adopted at an average temperature of 140 to 400°C, a hydrogen partial pressure of 1.0 to 10.0 MPa abs. and a space velocity per hour of 0.1 to 10.0 h ^-1 in order to obtain a hydrogenated effluent, b) after including at least one hydrogenation treatment The hydrotreating stage carried out in the hydrotreating reaction section of the catalyst, the hydrotreating reaction section is fed with at least the raw material or the hydrogenation effluent produced by stage a) and the gas stream containing hydrogen, the The hydrotreating reaction section is operated at an average temperature of 250 to 430°C, a hydrogen partial pressure of 1.0 to 10.0 MPa abs., and an hourly space velocity of 0.1 to 10.0 h ^-1 in order to obtain a hydrotreated effluent , c) a separation stage fed with the hydrotreated effluent produced by stage b) and optionally the hydrocracked effluent produced by stage g) and the aqueous solution, in order to obtain at least the gas effluent, the first aqueous effluent and the hydrocarbon effluent, d) the stage of separating the H ₂ S contained in the first aqueous effluent so as to obtain a gas phase containing H ₂ S and a second aqueous effluent, the The gas phase containing H ₂ S is recovered at least partially upstream of stage a) and/or stage b) and/or stage g), as appropriate, e) a stage of separation of NH ₃ contained in the second aqueous effluent , in order to obtain a gas phase containing NH ₃ and a third aqueous effluent, the gas phase containing NH ₃ being recovered at least partially upstream of stage a) and/or stage b) and/or stage g), as appropriate, f ) is the stage of optionally fractionating all or a part of the hydrocarbon effluent produced from stage c) in order to obtain at least a gaseous effluent and at least a first hydrocarbon fraction containing compounds having a boiling point less than or equal to 175°C and containing compounds having a boiling point greater than A second hydrocarbon fraction of compounds having a boiling point of 175°C, g) The hydrocracking stage is optionally carried out in a hydrocracking reaction zone using at least one fixed bed reactor with n catalytic beds, n being greater than or an integer equal to 1, each catalytic bed containing at least one hydrocracking catalyst, the hydrocracking reaction zone being fed at least a portion of the hydrocarbon effluent produced by stage c) and/or fed by stage f) At least a part of the second hydrocarbon fraction containing compounds having a boiling point greater than 175°C and a gas stream containing hydrogen produced, the hydrocracking reaction section is at an average temperature of 250 to 450°C, 1.5 to 20.0 MPa A hydrogen partial pressure of abs. and a space velocity per hour of 0.1 to 10.0 h ^-1 are used in order to obtain the first hydrocracked effluent.

The invention therefore relates to a method which makes it possible to purify the oil produced by the pyrolysis of plastic waste, which purifies at least a part of its impurities, which makes it possible to hydrogenate this oil and thus to be able to in particular by incorporating it directly into in a fuel storage unit or by making it compatible with processing in a steam cracking unit while enabling continuous recovery of the _H2S produced by the process in order to minimize the consumption of sulfiding agents. In particular at the beginning of the catalytic cycle (i.e. H ₂ S is formed for separation in stage d) and upstream of stage a) and/or stage b) and/or stage g), and/or also in the selective hydrogenation stage a0 ), the injection of vulcanizing agent is still required. Additional injections may be required throughout the catalytic cycle to compensate for natural losses. However, the fact that the gas phase containing H ₂ S but not NH ₃ can be recovered by the present invention makes it possible to significantly reduce the consumption of vulcanizing agents.

Another advantage is that NH ₃ in the gas effluent containing hydrogen and/or light hydrocarbons is removed from _the separation/washing section ( Stage c)) Top removal. In other words, the _NH3 remains in the aqueous effluent in the form of a salt.

Another advantage of the present invention is to prevent the risk of clogging and/or corrosion of the treatment unit in which the method of the present invention is carried out, which risks are exacerbated by the presence of diolefins, metals and halogen compounds in plastic pyrolysis oils, which are often present in large amounts.

The method of the present invention therefore makes it possible to obtain a hydrocarbon effluent produced from a plastic pyrolysis oil that at least partially releases the impurities of the starting plastic pyrolysis oil, thereby limiting the potential risks that may arise from such impurities. Operational problems, such as corrosion, coking or catalytic deactivation problems, particularly in the steam cracking unit and/or units located downstream of the steam cracking unit (in particular polymerization and hydrogenation units). The removal of at least part of the oil impurities produced by the pyrolysis of plastic waste will also allow the application range of the target polymer to be increased and use incompatibilities reduced.

According to an alternative form, the H ₂ S-containing gas phase produced by stage d) is recovered at least partially upstream of stage a) and/or stage b) and/or stage g).

According to an alternative form, the method includes a hydrogenation stage a).

According to an alternative form, the method includes a fractionation stage f).

According to an alternative form, the method includes a hydrocracking stage g).

According to an alternative, stage d) of separating the H ₂ S contained in the first aqueous effluent is by stripping the effluent with a steam-containing stream at a pressure of 0.5 to 1 MPa and a temperature of 80 to 150°C. to proceed.

According to an alternative, stage e) of separating the NH ₃ contained in the second aqueous effluent is carried out by stripping this effluent with a steam-containing stream at a pressure of 0.1 to 0.5 MPa and a temperature of 80 to 150°C. conduct.

According to an alternative form, the separation stage c) consists of the following stages: c1) a separation stage by feeding the hydrotreated effluent produced by stage b), which stage is carried out at a temperature between 200 and 450°C and at a pressure substantially equal to the pressure of stage b), in order to obtain at least gaseous effluent and liquid effluent, a part of which is recycled upstream of stage a) and/or stage b), as appropriate, c2) Separation stage, which is fed with the gaseous effluent produced by stage c1) and another part of the liquid effluent produced by stage c1) and an aqueous solution, this stage being at a temperature of 20 to less than 200°C and at a pressure substantially equal to or less than the pressure of stage b), so as to obtain at least a gas effluent, a first aqueous effluent and a hydrocarbon effluent.

According to an alternative form, the method includes at least one stage a0) of pre-treating a feedstock comprising plastic pyrolysis oil, optionally in a mixture with the hydrocarbon effluent produced in stage c), said pre-treatment stage being in stage a) Upstream and/or stage b) takes place upstream and includes a filtration stage and/or a centrifugation stage and/or an electrostatic separation stage and/or a stage of washing with the aid of an aqueous solution and/or an adsorption stage and/or a selective hydrogenation stage.

According to an alternative form, the hydrocarbon effluent produced by separation stage c) or at least one of the two liquid hydrocarbon fractions produced by stage f) is fed completely or partially to at least one pyrolysis furnace in Steam cracking stage h) carried out at a temperature of 700 to 900°C and a relative pressure of 0.05 to 0.3 MPa.

According to an alternative form, the NH ₃ -containing gas phase produced by stage e) is recovered at least partially upstream of stage a) and/or stage b) and/or stage g).

According to an alternative, a stream containing nitrogen compounds and/or sulfur compounds is injected upstream of stage a) and/or upstream of stage b).

According to an alternative form, the hydrogenation catalyst comprises a support selected from alumina, silica, silica-alumina, magnesium oxide, clays and mixtures thereof and at least one element from group VIII and at least one element from VIB Hydrogenation-dehydrogenation functionality of an element of group VIII, or of at least one element from group VIII.

According to an alternative form, the hydrotreating catalyst comprises a support selected from the group consisting of alumina, silica, silica-alumina, magnesium oxide, clays and mixtures thereof and at least one element from group VIII and/or hydrogenation-dehydrogenation functionality of at least one element from group VIB.

According to an alternative form, the process additionally comprises a second hydrocracking stage g') carried out in a hydrocracking reaction zone using at least one fixed bed reactor with n catalytic beds, n being greater than or equal to 1 an integer of , each catalytic bed containing at least one hydrocracking catalyst, the hydrocracking reaction zone being fed with at least a portion of the first hydrocracked effluent produced by the first hydrocracking stage g) A portion and a gas stream containing hydrogen, the hydrocracking reaction section is operated at a temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa abs., and an hourly space velocity of 0.1 to 10.0 h ^-1 , In order to obtain a second hydrocracked effluent.

According to an alternative form, the hydrocracking catalyst includes a support selected from the group consisting of halogenated alumina, boron and aluminum oxides, amorphous silica-alumina and zeolites and at least one selected from the group consisting of chromium, molybdenum and tungsten (alone or in mixtures), a Group VIB metal, and/or a hydrogenation-dehydrogenation function of at least one Group VIII metal selected from the group consisting of iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum.

The invention also relates to products readily obtainable by and preferably obtained according to the process according to the invention.

According to this alternative, the product contains, relative to the total weight of the product: - Metallic elements in a total content less than or equal to 10.0 ppm (by weight), -Contains iron at a level less than or equal to 200 ppb by weight, and/or - Silicon in an amount less than or equal to 5.0 ppm (by weight), and/or -Sulfur content less than or equal to 100 ppm (by weight), and/or - Nitrogen content less than or equal to 100 ppm (by weight), and/or -Chlorine content less than or equal to 10 ppm (by weight), and/or - Mercury content less than or equal to 5 ppb by weight.

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

According to the present invention, the expressions "of between ... and ..." and "between ... and ... .)" is equivalent and means that the limit of the interval is included in the stated range of values. If this is not the case and if the limits are not included within the stated range, this disclosure will be disclosed.

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

In the following, specific and/or preferred embodiments of the invention may be described. They can be used individually or in combination, with no combination restrictions when technically feasible.

Subsequently, the family of chemical elements is given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, editor-in-chief D.R. Lide, 81st edition, 2000-2001). For example, Group VIII (or VIIIB) according to the CAS classification corresponds to metals in lines 8, 9 and 10 according to the new IUPAC classification.

Metal content is determined by X-ray fluorescence.

raw material

According to the present invention, "plastic pyrolysis oil" is an oil that is advantageously in liquid form at ambient temperature and is produced from the heat of plastic, preferably plastic waste, in particular plastic waste gases originating from collection and sorting channels. solution produced. It can also be produced by the pyrolysis of worn tires.

They include in particular mixtures of hydrocarbon compounds, especially paraffins, mono- and/or diolefins, naphthalenes and aromatic hydrocarbons. At least 80% by weight of these hydrocarbon compounds preferably have a boiling point of less than 700°C and preferably less than 550°C. In particular, depending on the origin of the pyrolysis oil, the latter may contain up to 70% by weight of paraffins, up to 90% by weight of olefins and up to 90% by weight of aromatic hydrocarbons, it being understood that such paraffins, such The sum of olefins and these aromatic hydrocarbons is 100% by weight of hydrocarbon compounds.

The density of pyrolysis oil (measured at 15°C according to ASTM D4052 method) is generally between 0.75 and 0.99 g/cm ³ , preferably between 0.75 and 0.95 g/cm ³ .

The plastic pyrolysis oil may additionally and usually contains impurities such as metals, in particular iron, silicon or halogen compounds, in particular chlorine compounds. These impurities can be present in plastic pyrolysis oil at high levels, such as as high as 350 ppm (by weight) or 700 ppm (by weight), and actually even 1000 ppm (by weight) contributed by halogen compounds Halogen elements (especially chlorine), and up to 100 ppm (by weight), actually even 200 ppm (by weight) of metallic or semi-metallic elements. Alkali metals, alkaline earth metals, transition metals, late transition metals and metalloids can be placed in the same category as contaminants of metallic nature (called metals or metallic or semi-metallic elements). Specifically, the metals or metallic or semi-metallic elements that may be contained in the oil produced by the pyrolysis of plastic waste include silicon, iron or both of these elements. The plastic pyrolysis oil may also contain other impurities, such as impurities contributed in particular by sulfur compounds, oxygen compounds and/or nitrogen compounds, the content of which is generally less than 27 000 ppm (by weight) of impurities and preferably less than 15 500 ppm (by weight) of miscellaneous elements. These sulfides are generally present at a level of less than 2000 ppm (by weight) and preferably less than 500 ppm (by weight). These oxygen compounds are generally present in a content of less than 15 000 ppm (by weight) and preferably less than 10 000 ppm (by weight). These nitrogen compounds are generally present in amounts less than 10 000 ppm (by weight) and preferably less than 5000 ppm (by weight). The plastic pyrolysis oil may also contain other impurities such as heavy metals such as mercury, arsenic, zinc and lead, for example up to 100 ppb (by weight) or 200 ppb (by weight) of mercury.

The raw material for the method according to the invention contains at least one plastic pyrolysis oil. The raw material may consist solely of plastic pyrolysis oil. Preferably, the raw material contains at least 50% by weight, preferably 70% to 100% by weight of plastic pyrolysis oil relative to the total weight of the raw material, that is, preferably 50% to 100% by weight and in a relatively The best method is 70% to 100% by weight of plastic pyrolysis oil.

In addition to the plastic pyrolysis oil, the raw materials according to the method of the present invention may include conventional gasoline raw materials or raw materials produced by the conversion of biomass, which are then co-processed with the raw plastic pyrolysis oil.

It is known that the petroleum feedstock may advantageously be a fraction or a mixture of fractions of the naphtha, gas oil or vacuum gas oil type.

The raw materials resulting from the conversion of biomass may advantageously be selected from vegetable oils, oils from algae or seaweed oils, fish oils, waste food oils, and fats of vegetable or animal origin, or mixtures of such raw materials. Such vegetable oils may advantageously be wholly or partially crude or refined and are produced from the group consisting of oils from rape, sunflower, soybean, palm, olive, coconut, copra, castor oil plant, cotton plant, peanut oil, linseed oil and crambane oil. (sea kale oil), and all oils produced from plants such as sunflower or rapeseed through genetic modification or hybridization. This list is not limiting. The animal fats are advantageously chosen from blubber and fats consisting of residues from or due to the food industry. Frying oil, various animal oils, such as fish oil, beef tallow or lard can also be used. The raw material produced by the conversion of biomass may also advantageously be selected from methyl esters of fatty acids of plant and/or animal origin or of methyl esters of fatty acids of waste food vegetable oils.

The feedstock produced by the conversion of biomass can also be selected from feedstocks derived from methods for thermal or catalytic conversion of biomass, such as autogenous biomass, in particular from lignocellulosic biomass, using various liquefaction methods such as water Oil produced by thermal liquefaction or pyrolysis). The term "biomass" refers to material derived from recently living organisms, including plants, animals and their by-products. The term "lignocellulosic biomass" means biomass derived from plants or derived from by-products thereof. The lignocellulosic biomass is composed of carbohydrate polymers (cellulose, hemicellulose) and aromatic polymers (lignin).

The raw materials produced by the conversion of biomass may also advantageously be selected from those produced by the paper industry.

The plastic pyrolysis oil can be produced by thermal or catalytic pyrolysis treatment or prepared by hydropyrolysis (pyrolysis in the presence of a catalyst and hydrogen). Preprocessing ( optional )

The feedstock comprising plastic pyrolysis oil, optionally in a mixture with the hydrocarbon effluent from stage c), may advantageously be treated in at least one optional pretreatment stage a0) in the hydrogenation stage a) and/or in the hydrotreatment stage b ) to obtain pretreated raw materials for feeding stage a) and/or stage b).

According to an alternative form, this optional pretreatment stage a0) makes it possible to reduce the amount of contaminants and solid particles, in particular the amount of iron and/or silicon and/or chlorine that may be present in the feedstock containing the plastic pyrolysis oil. This optional stage a0) makes it possible in particular to remove sediments formed due to unstable properties of the pyrolysis oil and/or compatibility issues between the two different feedstocks. Therefore, it is advantageous to carry out the optional stage a0) of the pretreatment of the raw material containing plastic pyrolysis oil, especially when the raw material contains more than 10 ppm (by weight), especially more than 20 ppm (by weight), more particularly When it is more than 50 ppm (by weight) of metallic elements and/or solid particles, and especially when the raw material contains more than 5 ppm (by weight) of silicon, and more particularly more than 10 ppm (by weight), actual Even greater than 20 ppm (by weight) of silicon. Likewise, the optional stage a0) of the pretreatment of the raw material containing plastic pyrolysis oil is advantageously carried out, especially when the raw material contains more than 10 ppm (by weight), in particular more than 20 ppm (by weight), and more Especially when there is more than 50 ppm (by weight) chlorine.

This optional pretreatment stage a0) can be carried out by any method known to those skilled in the art which makes it possible to reduce the amount of contaminants. This may in particular include a filtration stage and/or a centrifugation stage and/or an electrostatic separation stage and/or a washing stage with the aid of an aqueous solution and/or an adsorption stage and/or a selective hydrogenation stage.

When the optional pretreatment stage a0) includes a filtration stage and/or a centrifugation stage and/or an electrostatic separation stage and/or a washing stage with the aid of an aqueous solution and/or an adsorption stage, advantageously at 20 to 400°C, preferably 40 to a temperature of 350°C and a pressure of 0.15 to 10.0 MPa abs., preferably 0.2 to 7.0 MPa abs.

According to an alternative form, the optional pretreatment stage a0) is carried out in an adsorption section operating in the presence of at least one adsorbent, preferably of the alumina type, having The specific surface area is greater than or equal to 100 m ² /g, preferably greater than or equal to 200 m ² /g. The specific surface area of the at least one adsorbent is advantageously less than or equal to 600 m ² /g, in particular less than or equal to 400 m ² /g. The specific surface area of the adsorbent is the surface area measured by the BET method, that is, by nitrogen adsorption according to the Brunauer-Emmett-Teller method described in The Journal of the American Chemical Society , 60, 309 (1938) The specific surface area measured according to the standard ASTM D 3663-78 established by the method.

Advantageously, the adsorbent contains less than 1% by weight of metal elements and is preferably free of metal elements. The term "metallic element of the adsorbent" should be understood as meaning the elements from rows 6 to 10 of the periodic table of the elements (new IUPAC classification). The residence time of the feedstock in the adsorption section is generally between 1 and 180 minutes.

The adsorption section of the optional stage a0) includes at least one adsorption tower, preferably at least two adsorption towers, preferably two to four adsorption towers, containing the adsorbent. When the adsorption section includes two adsorption towers, one mode of operation may be "swing" operation, where one of the towers is in-line, ie in operation, while the other tower is in reserve. When the adsorbent in the connected tower is exhausted, the tower is separated and the tower in reserve is placed online, that is, in operation. The spent adsorbent can then be regenerated in situ and/or replaced with fresh adsorbent so that the column containing it can be brought back online again when another column has been separated.

Another mode of operation is to have at least two columns operating in series. When the adsorbent in the tower placed at the head is exhausted, the first tower is separated and the exhausted adsorbent is regenerated in situ or replaced with a new adsorbent. The tower then brings back the connection at the last position and so on. This operation is called displacement mode, or according to the term PRS (Placable Reactor System), or "lead and lag". The combination of at least two adsorption towers makes it possible to overcome the adsorbent due to the combined action of metal contaminants, diolefins, gums produced from dienes and insoluble substances that may be present in the plastic pyrolysis oil to be treated and potential for rapid poisoning and/or clogging. The reason for this is that the presence of at least two adsorption towers facilitates the replacement and/or regeneration of the adsorbent, advantageously without the need to shut down the pretreatment unit, indeed even this method, thus making it possible to reduce the risk of blockage and thus avoid shutting down the unit due to blockage. , to control costs and limit adsorbent consumption.

According to another alternative, this optional pretreatment stage a0) is carried out in a section for washing with an aqueous solution (for example water) or an acidic or alkaline solution. The washing section may contain items of equipment which make it possible to contact the feedstock with the aqueous solution and to separate these phases in order to obtain on the one hand a pretreated feedstock and on the other hand to obtain an aqueous solution containing impurities. Such equipment items may include, for example, stirred reactors, decanters, mixer-decanters, and/or cocurrent or countercurrent scrubbers.

According to another alternative, the optional pretreatment stage a0) is performed by filtration. This filtration stage makes it possible to remove inorganic solids, sediments and/or fines, in particular metals, metal oxides and metal chlorides, contained in the feedstock. Filters are generally used whose pore size (eg diameter or equivalent diameter) is less than 25 µm, preferably less than or equal to 10 µm, and in an even better way less than or equal to 5 µm. According to another alternative, filters are used whose pore size is less than 25 µm but greater than 5 µm. It is also possible to use a series of filters with different pore sizes, in particular a series of filters with pore sizes that gradually decrease in the circulation direction of the raw material. Such filter media are well known from industrial applications. For example, cartridge filters or self-cleaning filters are suitable. Solids content may be determined, for example, by the Heptane Insolubles Test ASTM D-3279 method. The content of insoluble matter in heptane must be reduced to less than 0.5% by weight, preferably less than 0.1%.

According to a particular embodiment, the stage a0) of pretreatment by filtration includes at least one filter with a pore size of less than 10 µm and preferably greater than 5 µm, followed by a filtration system with a pore size of less than 2 µm And preferably less than 1 µm.

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

According to another specific embodiment, the stage a0) of pretreatment by filtration consists of at least one filter with a pore size of less than 10 µm and preferably greater than 5 µm, followed by the use of a filtration adjuvant such as sand or diatom soil) filter system.

According to another alternative, the optional pretreatment stage a0) is performed by centrifugation. According to another alternative, the pretreatment stage a0) includes centrifugation and filtration.

According to another alternative, the optional pretreatment stage a0) includes a selective hydrogenation stage. The selective hydrogenation stage is advantageously carried out in a reaction section feeding at least the feed and the gas stream containing hydrogen, optionally by one or more of the above-mentioned pretreatments, in the presence of at least one selective hydrogenation catalyst, at a temperature between 100 and The hydrogenation effluent is obtained at a temperature of 280°C, a hydrogen partial pressure of 1.0 to 10.0 MPa abs., and a space velocity per hour of 0.3 to 10.0 h ⁻¹ .

The selective hydrogenation stage is carried out under hydrogen pressure and temperature conditions such that the raw material can be maintained in the liquid phase and has a certain amount of soluble hydrogen. The soluble hydrogen is only necessary for the selective hydrogenation of dienes present in the pyrolysis oil. . This is advantageously carried out under milder conditions than in hydrogenation stage a). This selective hydrogenation of dienes in the liquid phase therefore makes it possible to avoid or at least limit the formation of "gums", that is, the polymerization of dienes and thus limit the formation of oligomers and polymers that can clog downstream reaction sections. Styrenic compounds (especially styrene) that may be present in the feedstock may also behave like dienes in terms of gum formation due to the fact that the vinyl double bonds are conjugated to the aromatic core. This selective hydrogenation stage makes it possible to obtain a selective hydrogenation effluent, that is to say an effluent with a reduced content of olefins, in particular diolefins and possibly styrenic compounds.

This reaction section uses selective hydrogenation, preferably in a fixed bed, in the presence of at least one selective hydrogenation catalyst, advantageously at 100 to 280°C, preferably 120 to 260°C, in a preferred way 130 to an average temperature of 250°C (or WABT as defined below for hydrogenation stage a)), a hydrogen partial pressure of 1.0 to 10.0 MPa abs., preferably 2.0 to 8.0 MPa abs., and a hydrogen partial pressure of 0.3 to 10.0 h ^-1 , preferably Preferably at an hourly space velocity (HSV) of 0.5 to 5.0 h ^-1 .

The amount of the gas stream containing hydrogen (H ₂ ) feeding the reaction section of the selective hydrogenation stage is advantageously such that the hydrogen coverage is from 1 to 200 Sm ³ hydrogen/m ³ feedstock (Sm ³ /m ³ ), preferably 1 To 50 Sm ³ hydrogen/m ³ feedstock (Sm ³ /m ³ ), in a preferred manner 5 to 20 Sm ³ hydrogen/m ³ feedstock (Sm ³ /m ³ ).

The hourly space velocity (HSV) and the hydrogen coverage are defined below for hydrogenation stage a).

This selective hydrogenation stage is preferably carried out in a fixed bed. It can also be carried out in an ebullated bed or in a moving bed.

Advantageously, the reaction section of the selective hydrogenation stage includes from 1 to 5 reactors. According to a specific embodiment of the invention, the reaction section includes 2 to 5 reactors, which operate in a replaceable mode (called PRS for a replaceable reactor system, or "forward" as described in hydrogenation stage a). and hysteresis"). According to a particularly preferred alternative form, the selective hydrogenation reaction section includes two reactors operating in displaceable mode.

The selective hydrogenation catalyst is generally a catalyst as described in hydrogenation stage a). It may be the same as or different from the catalyst of hydrogenation stage a).

The hydrogenation effluent obtained at the end of the selective hydrogenation stage has a reduced content of impurities, in particular dienes, relative to the content of the same impurities, in particular dienes, contained in the feedstock to the process. This selective hydrogenation stage generally allows conversion of at least 20% and preferably at least 30% of the dienes comprised in the starting feed.

Advantageously, this optional pretreatment stage a0) includes a filtration stage and/or a centrifugation stage and/or an electrostatic separation stage and/or a washing stage with the aid of an aqueous solution and/or an adsorption stage, followed by a selective hydrogenation stage.

The optional pretreatment stage a0) may optionally also be fed with the hydrocarbon effluent produced by stage c) of the process in the form of a mixture with the feedstock containing plastics pyrolysis oil or separately from the feedstock containing plastics pyrolysis oil. at least a part of the first hydrocarbon fraction produced by stage f) and/or a part of the first hydrocarbon fraction produced by stage f) comprising compounds having a boiling point less than or equal to 175°C and/or produced by stage f) comprising compounds having a boiling point greater than 175℃ part of the second hydrocarbon fraction. The recovery of at least a part of the liquid effluent produced by stage c) and/or of at least a part of the one or more hydrocarbon effluents produced by stage f) makes it possible to particularly increase the settling and thus improve the preparation of the feedstock after optional filtration. handle.

This optional pretreatment stage a0) thus makes it possible to obtain a pretreated starting material, which is subsequently fed in its presence to stage b) and/or to hydrogenation stage a). Hydrogenation stage a) ( optional )

According to the invention, the method optionally includes a hydrogenation stage a) carried out in a hydrogenation reaction zone using at least one fixed bed reactor with n catalytic beds, n being an integer greater than or equal to 1, each catalytic bed Containing at least one hydrogenation catalyst, the hydrogenation reaction section is fed with at least the optionally pretreated raw material and a gas stream containing hydrogen, and the hydrogenation reaction section is at an average temperature of 140 to 400°C, 1.0 to 10.0 MPa A hydrogen partial pressure of abs. and a space velocity per hour of 0.1 to 10.0 h ⁻¹ were used in order to obtain the hydrogenation effluent.

Stage a) is carried out in particular under conditions of temperature and hydrogen pressure which allow for the hydrogenation of the dienes which may remain after the optional selective hydrogenation stage and of the olefins at the beginning of the hydrogenation reaction section, and at the same time which allow the hydrogenation of the dienes to be carried out by the temperature An increasing curve for hydrodemetallization and hydrodechlorination, especially at the end of the hydrogenation reaction section. Injecting the necessary amount of hydrogen to make possible the hydrogenation of at least a part of the dienes and olefins present in the plastic pyrolysis oil, the hydrodemetallization of at least a part of the metals, in particular the retention of silicon, and also the at least part of the chlorine Partial conversion (to obtain HCl). The hydrogenation of dienes and olefins thus makes it possible to avoid or at least limit the formation of "gums", that is to say the polymerization of dienes and olefins and thus the formation of oligomers and polymerizations that can block the reaction section of the hydrotreatment stage b) things. In parallel with the hydrogenation, the hydrodemetallation and in particular the retention of silicon during stage a) make it possible to limit the catalytic deactivation of the reaction zone of the hydrotreatment stage b). Furthermore, the conditions of stage a) are such that at least part of the chlorine can be converted.

Temperature control is important in this stage and must satisfy antagonistic constraints. On the one hand, the inlet temperature and the temperature of the entire hydrogenation reaction zone must be high enough to allow hydrogenation of dienes and olefins at the beginning of the hydrogenation reaction zone. On the other hand, the inlet temperature of the hydrogenation reaction zone must be low enough to avoid deactivation of the catalyst. Since the hydrogenation reaction, especially the hydrogenation of part of the olefins and dienes, is highly exothermic, an increasing temperature profile is observed in the hydrogenation reaction zone. The higher temperature at the end of this section allows hydrodemetallization and hydrodechlorination reactions to proceed. Therefore, the outlet temperature of the reaction zone of stage a) is greater than the inlet temperature of the reaction zone of stage a), generally at least 3°C, preferably at least 5°C.

The temperature in stage a), whether it is the average temperature (WABT), the inlet temperature of the reaction zone, or the increase in temperature between the inlet and outlet of the reaction zone in stage a), can be determined in particular by diluent is injected into a), preferably a recycle of a part of the liquid effluent produced in stage c) and/or at least a part of the one or more hydrocarbon effluents produced in stage f), in particular by recycling ratio and/or controlled by the temperature of the recovered effluent.

The temperature difference between the inlet and the outlet of the reaction zone of stage a) is compatible with the injection of a gas (hydrogen) or a liquid cooling stream (especially part of the liquid hydrocarbon effluent produced by stage c).

The temperature difference between the inlet and the outlet of the reaction zone of stage a) is entirely due to the exothermic nature of the chemical reactions taking place in the reaction zone and therefore does not involve the use of heating means (oven, heat exchanger and the like).

The inlet temperature of the reaction zone of stage a) is between 135 and 397°C, preferably between 240 and 347°C.

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

According to the invention, the hydrogenation and the hydrogenation of dienes are advantageously carried out in one and the same stage and at a temperature sufficient to limit the deactivation of the catalyst of stage a) (which itself is manifested by a reduction in the conversion of dienes) part of the demetallization reaction. This same stage also makes it possible to benefit from the heat from the hydrogenation reaction, in particular the hydrogenation of part of the olefins and diolefins, in order to have an increasing temperature profile in this stage and thus to eliminate the need for catalytic additions located in the catalytic hydrogenation section. Requirements for heating devices between hydrogen treatment sections.

The reaction zone carries out hydrogenation in the presence of at least one hydrogenation catalyst, advantageously at an average temperature of 140 to 400°C, preferably 240 to 350°C and particularly preferably 260 to 330°C (or WABT as defined below), 1.0 to 10.0 MPa abs., preferably 1.5 to 8.0 MPa abs. hydrogen partial pressure and 0.1 to 10.0 h ^-1 , preferably 0.2 to 5.0 h ^-1 and excellent hourly space velocity of 0.3 to 3.0 h ^-1 ( HSV) under.

According to the present invention, the "average temperature" of the reaction zone corresponds to the weighted average bed temperature (WABT), which is well known to those skilled in the art. The average temperature is advantageously determined as a function of the catalytic system used, the equipment items, and the configuration thereof. This average temperature (or WABT) is calculated as follows: Among them, T _inlet is the temperature of the effluent at the inlet of the reaction section, and T _outlet is the temperature of the effluent at the outlet of the reaction section. Unless otherwise indicated, the "average temperature" of the reaction zone is given at the start of the cycle.

Hourly space velocity (HSV) is defined here as the ratio of the hourly flow rate (in volume) of the optionally pretreated feedstock containing plastic pyrolysis oil to the volume of the catalyst.

The hydrogen coverage is defined as the ratio of the flow rate (in volume) of hydrogen used under standard temperature and pressure conditions relative to the flow rate (in volume) of the "fresh" feedstock (that is, the feedstock to be processed), which Pretreatment at 15°C has been carried out as appropriate, without taking into account the recovered fractions and in particular without taking into account the recovered liquid effluent resulting from stage c) and/or the liquid effluent resulting from stage f) (The unit is standard m ³ , indicating Sm ³ of H ₂ per m ³ of raw material).

The amount of the gas stream containing hydrogen (H ₂ ) feeding the reaction section of stage a) is advantageously such that the hydrogen coverage is from 100 to 1500 Sm ³ hydrogen/m ³ feedstock (Sm ³ /m ³ ), preferably from 200 to 200 1000 Sm ³ hydrogen/m ³ feedstock (Sm ³ /m ³ ), in a preferred manner 250 to 800 Sm ³ hydrogen/m ³ feedstock (Sm ³ /m ³ ).

The hydrogen can be produced from fossil sources or from renewable sources, such as from the gasification of plastic waste, or by electrolysis.

Advantageously, the reaction section of stage a) includes 1 to 5 reactors, preferably 2 to 5 reactors, and particularly preferably it includes two reactors. The advantage of a hydrogenation reaction section containing several reactors lies in the optimal treatment of the feedstock and at the same time makes it possible to reduce the risk of clogging of the catalytic bed and thus avoid shutting down the unit due to clogging.

According to a preferred alternative, the reactors are operated in displaceable mode (called PRS, or "lead and lag" for displaceable reactor systems). The combination of at least two reactors in PRS mode makes it possible to separate the reactors, drain the spent catalyst, refill the reactors with fresh catalyst and put the reactor back into service without shutting down the process. This PRS technology is described in particular in patent FR 2 681 871.

According to a particularly preferred alternative, the hydrogenation reaction section of stage a) consists of two reactors operated in displaceable mode.

Advantageously, reactor internals (eg of the filter plate type) can be used to prevent clogging of the reactor. An example of a filter plate is described in patent FR 3 051 375.

Advantageously, the hydrogenation catalyst contains a support (preferably an inorganic support) and hydrogenation-dehydrogenation functionality.

According to an alternative form, the hydrogenation-dehydrogenation function includes in particular at least one element from group VIII, preferably selected from nickel and cobalt, and at least one element from group VIB, preferably selected from molybdenum and tungsten. According to this alternative form, the total content of metal elements from Group VIB and Group VIII (expressed as oxides) is preferably between 1% and 40% by weight, preferably 5% by weight, relative to the total weight of the catalyst. % to 30% by weight. When the metal is cobalt or nickel, the metal content is expressed as CoO and NiO respectively. When the metal is molybdenum or tungsten, the metal content is expressed as MoO ₃ and WO ₃ respectively.

The weight ratio expressed as metal oxide of the Group VIB metal (or metals) relative to the Group VIII metal (or metals) is preferably between 1 and 20 and in a preferred manner between 2 and 10.

According to this alternative, the reaction section of stage a) includes, for example, 0.5 to 12% by weight of nickel, preferably 0.9 to 10% by weight of nickel (relative to the weight of the catalyst, in terms of The hydrogenation catalyst is nickel oxide NiO) and 1 to 30 wt% molybdenum, preferably 3 to 20 wt% molybdenum (relative to the weight of the catalyst, expressed as molybdenum oxide MoO ₃ ). Preferably on an inorganic support, preferably on an alumina support.

According to another alternative, the hydrogenation-dehydrogenation function includes and preferably consists of at least one element from group VIII, preferably nickel. According to this alternative form, the content of nickel oxide is preferably between 1% and 50% by weight, preferably between 10% and 30% by weight relative to the weight of the catalyst. Catalysts of this type are preferably used in their reduced form, preferably on an inorganic support, preferably on an alumina support.

The carrier of the hydrogenation catalyst is preferably selected from the group consisting of alumina, silica, silica-alumina, magnesium oxide, clay and mixtures thereof. The support may comprise a dopant compound, in particular an oxide selected from the group consisting of boron oxide (especially boron trioxide), zirconium oxide, ceria, titanium oxide, phosphorus pentoxide and mixtures of these oxides. Preferably, the hydrogenation catalyst comprises an alumina support optionally doped with phosphorus and optionally boron. When phosphorus pentoxide _P2O5 is present, its concentration is less than 10% by weight relative to the weight of alumina, and advantageously at least 0.001% by weight relative to the total _weight of alumina. When diboron trioxide _B2O3 is present, its concentration is less than 10% by weight relative to the weight of alumina, and advantageously _at least 0.001% relative to the total weight of alumina. The aluminum oxide used may be, for example, a gamma (gamma) or an eta (eta) alumina.

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

Advantageously, stage a) may employ, in addition to the hydrogenation catalyst described above, at least one hydrogenation catalyst on an alumina support used in stage a), which contains less than 1% by weight relative to the weight of the catalyst. Nickel and at least 0.1% by weight of nickel, preferably 0.5% by weight of nickel (expressed as nickel oxide NiO), and less than 5% by weight of molybdenum and at least 0.1% by weight of molybdenum relative to the weight of the catalyst, preferably 0.5% by weight of molybdenum (expressed as molybdenum oxide, MoO ₃ ). The catalyst (not highly loaded with metal) can be preferably placed upstream or downstream (preferably upstream) of the above-mentioned hydrogenation catalyst.

Preferably, stage a) can use at least one guard bed upstream of the hydrogenation catalyst, which contains alumina type, silica-alumina type, zeolite type and/or activated carbon type adsorbent, which optionally contains from Metals of Group VIB and/or Group VIII. It is also possible to use a series of beds with particles of different diameters, in particular a series of beds with decreasing diameters in the direction of circulation of the raw material (also called "grading").

This hydrogenation stage a) makes it possible to obtain a hydrogenation effluent, that is to say an effluent having a reduced content of olefins (especially dienes) and metals (especially silicon). The hydrogenation effluent obtained at the end of stage a) has a reduced content of impurities, in particular dienes, relative to the content of the same impurities, in particular dienes, contained in the feedstock to the process. The hydrogenation stage a) generally results in the conversion of at least 40% and preferably at least 60% of the dienes and at least 40% and preferably at least 60% of the olefins comprised in the starting feed. The heat released by the saturation of the double bonds makes it possible to increase the temperature of the reaction medium and initiate the hydrotreating reaction, in particular the at least partial removal of other contaminants such as, for example, silicon and chlorine. Preferably, at least 50% and more preferably at least 75% of the chlorine and silicon of the starting feedstock are removed during stage a). The hydrogenation effluent obtained at the end of hydrogenation stage a) is preferably sent directly to hydrotreatment stage b). Hydrotreating stage b)

According to the invention, the treatment method includes a hydrotreating stage b) carried out in a hydrotreating reaction zone containing at least one hydrotreating catalyst fed at least optionally in stage a0 ) or the hydrogenation effluent and gas stream containing hydrogen produced in stage a), the hydrogenation reaction section is hydrogen at an average temperature of 250 to 430°C and 1.0 to 10.0 MPa abs. Partial pressures and space velocities per hour of 0.1 to 10.0 h ^-1 are used in order to obtain a hydrotreated effluent.

Advantageously, stage b) carries out hydrotreating reactions well known to those skilled in the art and more particularly hydrotreating reactions, such as hydrogenation, hydrodesulfurization and hydrodenitrification of aromatic hydrocarbons. In addition, hydrogenation and hydrodemetallization of olefins and remaining halogen compounds continue.

The hydrotreating reaction zone is advantageously at a pressure equivalent to the pressure in the reaction zone for hydrogenation stage a), and generally at a higher average temperature than the average temperature of the reaction zone of hydrogenation stage a) Adopted below. Therefore, the hydrotreating reaction section is advantageously at an average hydrotreating temperature of 250 to 430°C, preferably 280 to 380°C, a hydrogen partial pressure of 1.0 to 10.0 MPa abs., and 0.1 to 10.0 h ^-1 , 0.1 to 5.0 h ^-1 , preferably 0.2 to 2.0 h ^-1 , and in a preferred manner implemented at an hourly space velocity (HSV) of 0.2 to 1 h ^-1 . The hydrogen coverage in stage b) is advantageously 100 to 1500 Sm ^hydrogen / ^m fresh feed, preferably 200 to 1000 Sm ^hydrogen / ^m fresh feed and in a preferred manner 250 to ⁸⁰⁰ Sm Hydrogen/m ³ new raw materials. The definitions of average temperature (WABT), HSV and hydrogen coverage correspond to those described above.

This hydrotreating stage is preferably carried out in a fixed bed. It can also be carried out in ebullated beds, entrained beds or moving beds. When the hydrotreatment stage is carried out in an ebullated bed, an entrained bed or a moving bed, another stage of hydrotreatment in a fixed bed may be carried out within the same range of operating conditions, followed by an ebullated bed, an entrained bed or a moving bed. In a moving bed, it is carried out with or without an intermediate separated gas stream.

Preferably, the treatment method includes a hydrotreating stage b) carried out in a hydrotreating reaction section, which uses at least one fixed bed reactor with n catalytic beds, where n is an integer greater than or equal to 1, each The catalytic bed contains at least one hydrotreating catalyst.

The hydrotreating reaction section is advantageously operated at the first catalytic bed of the first reactor by being fed at least the feedstock or the hydrogenation effluent produced from stage a) and the gas stream comprising hydrogen. It is also possible to inject at least a portion of the feedstock or the hydrogenation effluent produced from stage a) and/or at least a portion of the hydrogen between the different catalytic beds. Optionally, the reaction section of stage b) may also be fed in addition at least a portion of the liquid effluent produced by stage c) and/or at least a portion of one of the effluents from stage f).

Advantageously, this stage b) is carried out in a hydrotreating reaction section comprising at least one, preferably one to five fixed bed reactors with n catalytic beds, n being greater than or equal to 1, preferably 1 to 10. In a preferred manner, an integer from 2 to 5, each of the beds includes at least one and preferably no more than ten hydrotreating catalysts. When the reactor includes several catalytic beds, that is at least two, preferably two to ten, and in a preferred way two to five catalytic beds, these catalytic beds are preferably arranged in series in the reaction in the vessel.

When stage b) is carried out in a hydrotreating reaction section comprising several reactors, preferably two reactors, these reactors can be connected in series and/or in parallel and/or in replaceable (or PRS) mode and /or operate in swing mode. The various selectable operating modes, PRS (or lead and lag) modes and swing modes are well known to those skilled in the art and are advantageously defined above.

In another embodiment of the present invention, the hydrotreating reaction section includes a single fixed bed reactor containing n catalytic beds, n is greater than or equal to 1, preferably between 1 and 10, preferably between 2 An integer between 5 and 5.

In a particularly preferred embodiment, the hydrogenation reaction section of stage a) includes two reactors operating in a replaceable mode, followed by the hydrotreating reaction section of stage b), which includes a single fixed bed reactor .

Advantageously, the hydrotreating catalyst used in this stage b) can be chosen from known hydrodemetallization, hydrotreating or silicon scavenging catalysts used in particular for the treatment of petroleum fractions, and combinations thereof. Hydrodemetallation catalysts are known, for example, as described in patents EP 0 113 297, EP 0 113 284, US 5 221 656, US 5 827 421, US 7 119 045, US 5 622 616 and US 5 089 463. By. Known hydrotreating catalysts are, for example, those described in patents EP 0 113 297, EP 0 113 284, US 6 589 908, US 4 818 743 or US 6 332 976. Known silicon removal catalysts are, for example, those described in patent applications CN 102051202 and US 2007/080099.

In particular, the hydrotreating catalyst includes a support (preferably an inorganic support) and at least one metal element with hydrogenation-dehydrogenation functionality. The metal element with hydrogenation-dehydrogenation functionality advantageously contains at least one element from group VIII, preferably selected from the group consisting of nickel and cobalt, and/or at least one element from group VIB, preferably selected from the group consisting of molybdenum and tungsten. of elements. Relative to the total weight of the catalyst, the total content (expressed as oxides) of metal elements from Group VIB and Group VIII is preferably between 0.1% and 40% by weight, preferably between 35% and 30% by weight. . When the metal is cobalt or nickel, the metal content is expressed as CoO and NiO respectively. When the metal is molybdenum or tungsten, the metal content is expressed as MoO ₃ and WO ₃ respectively. The weight ratio expressed as metal oxide of the Group VIB metal relative to the Group VIII metal is preferably between 1.0 and 20, in a preferred manner between 2.0 and 10. For example, the hydrotreating reaction section of stage b) of the method includes a hydrotreating catalyst on an inorganic support, preferably on an alumina support, which includes a weight relative to the total weight of the hydrotreating catalyst. 0.5% to 10% by weight of nickel, preferably 1% to 8% by weight of nickel (expressed as nickel oxide NiO), and 1.0% to 30% by weight relative to the total weight of the hydrotreating catalyst. Molybdenum, preferably 3.0% to 29% by weight of molybdenum (expressed as molybdenum oxide MoO ₃ ).

The support for the hydrotreating catalyst is advantageously selected from the group consisting of alumina, silica, silica-alumina, magnesium oxide, clays and mixtures thereof. The support may additionally comprise a dopant compound, in particular an oxide selected from the group consisting of boron oxide (especially boron trioxide), zirconium oxide, ceria, titanium oxide, phosphorus pentoxide and mixtures of these oxides. Preferably, the hydrotreating catalyst comprises an alumina support, preferably an alumina support doped with phosphorus and optionally boron. When phosphorus pentoxide _P2O5 is present, its concentration is less than 10% by weight relative to the weight of alumina, and advantageously at least 0.001% by weight relative to the total _weight of alumina. When diboron trioxide _B2O3 is present, its concentration is less than 10% by weight relative to the weight of alumina, and advantageously _at least 0.001% relative to the total weight of alumina. The aluminum oxide used may be, for example, a gamma (gamma) or an eta (eta) alumina.

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

Advantageously, the hydrotreating catalyst used in stage b) of the process exhibits a specific surface area greater than or equal to 250 m ² /g, preferably greater than or equal to 300 m ² /g. The specific surface area of the hydrotreating catalyst is advantageously less than or equal to 800 m ² /g, preferably less than or equal to 600 m ² /g, especially less than or equal to 400 m ² /g. The specific surface area of the hydrotreating catalyst is the surface area measured by the BET method, that is, by nitrogen adsorption according to Brunauer-Emmett described in The Journal of the American Chemical Society , 60, 309 (1938) -Measured according to the standard ASTM D 3663-78 established by Teller method. This specific surface area makes it possible to further improve the removal of contaminants, especially metals such as silicon.

According to another aspect of the present invention, the above-mentioned hydrotreating catalyst additionally contains one or more organic compounds containing oxygen and/or nitrogen and/or sulfur. This catalyst is often referred to by the term "additive catalyst". Generally speaking, the organic compound is selected from the group consisting of one or more carboxylic acids, alcohols, thiols, thioethers, sulfates, sulfoxides, ethers, aldehydes, ketones, esters, carbonates, amines, nitriles, acylases Chemically functional compounds with amine, oxime, urea and amide functions may also include furan rings or sugar compounds.

In a preferred way, stage b) can use at least one guard bed or a series of guard beds of the "staged" type as described above for stage a) upstream of the hydrogenation catalyst.

Advantageously, this hydrotreatment stage b) allows at least 80% and preferably all of the olefins remaining after hydrogenation stage a) to be hydrogenated and to at least partially convert other impurities present in the feed (such as aromatic compounds, metal compounds, Sulfur compounds, nitrogen compounds, halogen compounds (especially chlorine compounds) and oxygen compounds). Preferably, the nitrogen content at the outlet of stage b) is less than 100 ppm (by weight). Preferably, the sulfur content at the outlet of stage b) is less than 100 ppm (by weight). Stage b) may also make it possible to further reduce the content of contaminants, such as the content of metals, especially silicon. Preferably, the metal content at the outlet of stage b) is less than 10 ppm (by weight), in a preferred way less than 2 ppm (by weight), and the silicon content is less than 5 ppm (by weight) ).

Depending on the content of sulfur compounds in the initial feedstock to be treated, the stream containing the vulcanizing agent can be produced upstream of the optional pretreatment stage a0), the optional hydrogenation stage a) and/or the hydrotreating stage b) and/or in the optional Injection upstream of one of the hydrocracking stages g), when present, preferably upstream of the hydrogenation stage a) and/or the hydrotreating stage b), in order to ensure a sufficient amount of sulfur forming catalyst (in the form of sulfide) ) active entity.

This activation or sulfidation stage is carried out by methods well known to those skilled in the art and advantageously in 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 fractions containing sulfur compounds with a boiling point of less than 400°C or used to focus on sulfurization contacts. Any other sulfur-containing compound that acts as a mediator to activate hydrocarbon feedstocks. Such sulfur-containing compounds are advantageously selected 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-butane mercaptan), and polysulfide compounds of the third nonyl polysulfide type. The catalyst can also be vulcanized by sulfur contained in the raw material to be desulfurized. Preferably, the catalyst is vulcanized in situ in the presence of a vulcanizing agent and hydrocarbon feedstock. Advantageously, the catalyst is vulcanized in situ in the presence of feedstock to which dimethyl disulfide has been added.

At the beginning of the catalytic cycle (i.e. H ₂ S is formed for separation in stage d) and upstream of stage a) and/or stage b) and/or stage g) or upstream of the selective hydrogenation stage of pretreatment a0) time for recycling), especially the injection of vulcanizing agent. Additional injections may be required throughout the catalytic cycle to compensate for natural losses. However, the fact that the gas phase containing H ₂ S but not NH ₃ can be recovered by the present invention makes it possible to significantly reduce the consumption of vulcanizing agents. Separation stage c)

According to the invention, the treatment method comprises a separation stage c), advantageously carried out in at least one washing/separation section, which is fed at least the hydrotreated effluent produced by stage b) and optionally The hydrocracked effluent and the aqueous solution produced by optional stages g) and g') are introduced in order to obtain at least a gas effluent, a first aqueous effluent and a hydrocarbon effluent.

The gas effluent obtained at the end of stage c) advantageously contains hydrogen, preferably at least 80% by volume, preferably at least 85% by volume hydrogen. Advantageously, the gas effluent may be at least partially recycled to the hydrogenation stage a) and/or the hydrotreating stage b) and/or the hydrocracking stage g), the recovery system may comprise a purification section.

This gas effluent may also constitute the subject of a further separation for the purpose of recovering at least one hydrogen-rich gas and/or light hydrocarbons, in particular ethane, propane and butane, which may be advantageously used alone or in a mixture Sent to one or more furnaces of steam cracking stage h) in order to increase the overall yield of olefins.

The hydrocarbon effluent produced by separation stage c) is fed partially or completely directly to the inlet of the steam cracking unit or to optional fractionation stage f). Preferably, the liquid hydrocarbon effluent is sent partially or completely to fractionation stage f).

The first aqueous effluent obtained at the end of stage c) advantageously contains ammonium salts and/or hydrochloric acid, and also dissolved _H2S and _NH3 .

This separation stage c) makes it possible to specifically remove ammonium chloride salts, which are formed by the reaction between chloride ions and ammonium ions, which chloride ions are released in the form of HCl by the hydrogenation of the chlorine compound, especially in stages a) and During b), the ammonium ions are produced in the form of _NH3 by hydrogenation of nitrogen compounds, especially during stage b), and/or contributed by the injection of amines, then dissolved in water, and are therefore limited due to Risk of clogging due to precipitation of ammonium chloride salts, in particular in transfer lines and/or in sections of the process according to the invention and/or in transfer lines to the steam cracker. It also allows the removal of hydrochloric acid formed by the reaction of hydrogen ions and chloride ions. This stage c) also allows the removal of ammonium sulfide salts ((NH ₄ ) ₂ S), which are formed by the reaction between H ₂ S and NH ₃ produced by hydrodesulfurization of sulfur compounds.

As a function of the content of chlorine compounds in the starting feedstock to be treated, streams containing nitrogen compounds such as amines, for example monoethanolamine, diethanolamine and/or monodiethanolamine can be produced upstream of the selective hydrogenation stage of pretreatment a0) and/or Injection upstream of hydrogenation stage a) and/or between hydrogenation stage a) and hydrotreating stage b) and/or between hydrocracking stage g) and separation stage c), preferably upstream of hydrogenation stage a), In order to ensure a sufficient amount of ammonium ions combined with the chloride ions formed during the hydrotreating stage, thus making it possible to limit the formation of hydrochloric acid and thus corrosion downstream of the separation section.

According to an alternative form, the NH ₃ -containing gas phase produced by stage e) can also be used as nitrogen compound.

Advantageously, the separation stage c) includes, upstream of the washing/separation section, the injection of an aqueous solution, preferably water, into the hydrogenation effluent produced by stage b) or, alternatively, by the added hydrogenation effluent produced by stage g). in the effluent of hydrogen cracking in order to at least partially dissolve ammonium chloride salts and/or hydrochloric acid and thereby improve the removal of chlorinated impurities and reduce the risk of clogging due to accumulation of ammonium chloride salts.

This separation stage c) is advantageously carried out at a temperature of 20 to 450°C, preferably 100 to 440°C, preferably 200 to 420°C. It is important to operate within this temperature range (and thus not overcool the hydrotreated effluent) without the risk of blockage in the line due to precipitation of ammonium chloride salts. Advantageously, this separation stage c) is carried out at a pressure close to that used in stages a) and/or b), preferably at a pressure between 1.0 and 10.0 MPa, in order to facilitate the recovery of hydrogen.

This separation stage may advantageously be achieved by any method known to those skilled in the art, such as for example a combination of one or more separators (drums) and/or one or more stripping columns. To proceed, the separator (drum) and/or column(s) may optionally be fed with a stripping gas (eg a hydrogen-rich gas stream). The washing/separating section of stage c) may be made at least partly from washing and separation equipment jointly or separately.

In an alternative embodiment of the invention, the separation stage c) consists of injecting an aqueous solution into the hydrotreated effluent produced by stage b), followed by a washing/separation section advantageously comprising a separation phase , which makes it possible to obtain at least one first aqueous effluent to which ammonium salt is added, a washed liquid hydrocarbon effluent and a partially washed gaseous effluent. The ammonium salt-added first aqueous effluent and the washed liquid hydrocarbon effluent may be subsequently separated in a knockout drum to obtain the hydrocarbon effluent and the first aqueous effluent. The partially scrubbed gas effluent may be introduced in parallel into a scrubber column where it is recycled countercurrently to an aqueous stream preferably having the same properties as the aqueous solution injected into the hydrotreated effluent, This makes it possible to remove at least partially, preferably completely, the hydrochloric acid contained in the partially scrubbed gas effluent and thus obtain the gas effluent and the acidic aqueous stream, preferably essentially containing hydrogen. The first aqueous effluent produced by the separation drum is optionally mixed with the acidic aqueous stream and optionally used in a mixture with the acidic aqueous stream in a water recovery loop for separation upstream of the wash/separation section Stage c) feeds the aqueous solution and/or feeds the aqueous stream in the scrubber. The water recovery loop may include a supply and/or bleed of water and/or alkaline solution so that dissolved salts can be drained off.

In another alternative embodiment of the invention, the separation stage c) may advantageously comprise a "high pressure" washing/separation section, which may occur close to the hydrogenation stage a) and/or the hydrotreating stage b) and/or may The pressure of the hydrocracking stage g) is selected to operate at a pressure (preferably between 1.0 and 10.0 MPa) in order to promote the recovery of hydrogen. This optional "high-pressure" section of stage c) may be supplemented by a "low-pressure" section in order to obtain a gas fraction free from dissolution at high pressure and intended to be processed directly in a steam cracking process or optionally sent to fractionation stage f) of hydrocarbon effluents. Variant two-stage separation stage c)

In an alternative embodiment of the invention, the separation stage c) includes the following sub-stages: c1) a separation stage fed with the hydrotreated effluent produced by stage b), which stage is at 200 to a temperature of 450°C and at a pressure substantially equal to the pressure of stage b), so as to obtain at least a gaseous effluent and a liquid effluent, part of which can be recovered upstream of stage a) and/or stage b), c2 ) a separation stage, which is fed the gas effluent produced by stage c1) and another part of the liquid effluent produced by stage c1) and an aqueous solution, which stage is at a temperature of 20 to less than 200°C and This is carried out at a pressure substantially equal to or less than the pressure of stage b), so as to obtain at least a gas effluent, a first aqueous effluent and a hydrocarbon effluent. Stage c1)

According to the invention, the treatment method may comprise a separation stage c1) fed with the hydrotreated effluent produced by stage b) at a temperature between 200 and 450°C and substantially equal to the stage It is carried out under the pressure of b) in order to obtain at least gas effluent and liquid effluent, part of which can be recovered upstream of stage a) and/or stage b).

This separation stage c1) is what is known as the high pressure or medium pressure high temperature separation stage, also known by the name HHPS (Hot High Pressure Separator) to those skilled in the art. Therefore, it is better to use a "hot high-pressure" separator in this stage c1), and the pressure is essentially equal to the operating pressure of stage b). The term "substantially equal to the pressure of stage b)" is to be understood as meaning a pressure of stage b) whose pressure difference relative to the pressure of stage b) is 0 to 1 MPa, preferably 0.005 to 0.3 MPa and particularly preferably 0.01 to 0.01 MPa. 0.2 MPa. Preferably, the pressure in stage c1) is the pressure reduced due to the pressure drop in stage b).

The temperature at which the separation is carried out is between 200 and 450°C, preferably between 220 and 330°C and particularly preferably between 240 and 300°C. According to a better alternative and with an eye on recovering the majority of the heat, the separation is carried out at the highest possible temperature but less than or equal to the outlet temperature of stage b), which makes it possible to avoid or limit the reheating of the effluent from stage b) (and therefore requires heat). According to another alternative, the effluent from stage b) can be reheated or cooled before separation.

This separation stage c1) may advantageously be carried out by any method known to those skilled in the art, such as for example a combination of one or more separators (drums) and/or one or more stripping columns, The separator (drum) and/or column(s) may optionally be fed with a stripping gas (eg a hydrogen-rich gas stream). Preferably, stage c) is carried out using a single separator (drum).

A part of the liquid effluent can be recovered upstream of stage a) and/or stage b) and/or upstream of the selective hydrogenation stage a0) of the pretreatment. Recycling a part of the product obtained to at least one of the reaction stages or upstream thereof advantageously makes it possible to dilute the impurities on the one hand and to control the temperature in the reaction stage on the other hand, wherein the reactions involved may be Highly exothermic.

Advantageously, the amount of the recovered liquid effluent produced by stage c) (ie the recovered fraction of the product obtained) is adjusted so that the recovered stream from stage c) is identical to the feedstock comprising plastic pyrolysis oil ( That is, the ratio (by weight) of the raw material to be processed feeding the overall method 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 in a preferred The method is greater than or equal to 0.1. Preferably, the amount of recovered liquid effluent produced by stage c) is adjusted such that the ratio (by weight) of the recovered stream to the feedstock containing the plastic pyrolysis oil is between 0.01 and 10, preferably between 0.1 to 7 and particularly preferably between 0.2 and 5. This recovery ratio makes it possible to specifically control the temperature rise in stage a). This is because, when the recovery ratio is high, the dilution rate of the feedstock is high and the rise in temperature at the beginning of the reaction section of stage a), in particular due to the hydrogenation of dienes, can therefore be controlled by the dilution effect.

This high-pressure and high-temperature separation makes it possible to maximize energy recovery by thermally recovering a portion of the liquid effluent on the one hand. This is because the energy required to reach the necessary inlet temperatures in stage a) and/or stage b) can be at least partially contributed by the heat of a portion of the liquid effluent produced by stage c) and can be reduced, practically even Optional preheating by direct heating of raw materials beyond temperatures greater than 200°C is dispensed with to prevent glue formation. Furthermore, the fact that at least part of the liquid effluent is recovered at high pressure makes it possible to save energy for its pressurization in stage a) and/or stage b).

This high-pressure and high-temperature separation makes it possible, on the other hand, to minimize the light fractions (hydrocarbons containing compounds with a boiling point less than or equal to 175°C) contained in the liquid effluent recovered in stage a) and/or stage b) distillate or naphtha). At this temperature, virtually all the light fractions of the effluent (naphtha) remain as gas effluent in the separation/washing stage c2), while as the liquid phase mainly the heavy fractions of the feedstock (including those with greater than 175 Hydrocarbon fractions or middle distillates of compounds with boiling points of ℃). In this way, this pH ₂ p is favorable for stage a) and/or stage b), since the light fraction (naphtha) can partially evaporate and reduce if it is not removed at least partially during the high-pressure and high-temperature separation pH _2p . The removal of the light fractions containing naphtha may optionally be increased by a slight pressure reduction upstream of the at least one separator used in stage c1), even if this use is not preferred due to the energy losses associated with the pressure reduction. Another option for increasing the removal of light ends containing naphtha may include performing stripping, for example by injecting a hydrogen-rich gas in stage c1).

According to a preferred alternative, at least a part of the hydrotreated liquid effluent produced from stage c) can advantageously be cooled or preheated, if necessary, or advantageously during the hydrogenation stage a) and/or Before recycling upstream of the hydrogen treatment stage b), the temperature and flow rate of the feedstock and hydrogen are maintained at the same temperature as at the outlet of the separation stage c), so that the liquid effluent produced by stage c) is contained. The temperature of the feed stream of at least a part of the mixture and the incoming stream of hydrogen-rich gas is between 140 and 430°C, preferably between 220 and 350°C and particularly preferably between 260 and 330°C.

In the case where at least a portion of the liquid effluent produced by separation stage c) is preheated before recovery upstream of stage a) and/or stage b), such effluent is treated in stage a) and/or stage b) as appropriate. b) Passing through at least one exchanger and/or at least one oven prior to upstream recovery in order to adjust the temperature of the recovered liquid effluent.

In the case where at least a portion of the liquid effluent produced by separation stage c) is cooled before being recycled upstream of stage a) and/or stage b), such effluent is treated in stage a) and/or stage b as appropriate. ) is passed through at least one exchanger and/or at least one cooling tower prior to upstream recovery in order to adjust the temperature of the recovered liquid effluent.

Thus, at least a portion of the liquid effluent produced from stage c) upstream of the hydrogenation stage a) and/or upstream of the hydrotreating stage b) (which may be cooled or preheated as required, or maintained in the separation stage The use of recycle material with the same temperature at the outlet of c) allows the temperature of the stream entering stage a) and/or stage b) to be adjusted as needed.

According to an alternative form, the feedstock can be heated by direct heating to a temperature of up to 200°C, preferably up to 180°C and particularly preferably up to 150°C, before being mixed with at least a part of the effluent produced by stage c). Preheating is performed within the temperature range. Above this temperature, contact with the walls during direct heating can lead to the formation of gum and/or coke, which can lead to fouling and an increase in the pressure drop in the system used to heat the feedstock and also in the catalyst bed. Heating the feedstock to a temperature above 150°C, preferably above 180°C and particularly preferably above 200°C is preferably carried out by indirect heating of at least part of the effluent produced by stage c). Therefore, an increase in the temperature of the raw material above 150°C, preferably above 180°C and particularly preferably above 200°C is achieved by mixing with the hotter liquid rather than by contact with the heated wall. According to another alternative, the raw material is heated to a temperature above 150°C, preferably above 180°C and particularly preferably above 200°C by means of a proportional oven or exchanger in order to have a very low wall compared to the temperature of the raw material temperature (such as an electric oven).

According to another alternative, the feedstock is completely heated by indirect heating of at least a portion of the effluent produced by stage c). In this case, the feedstock is not preheated before being mixed with at least a part of the effluent produced by stage c).

According to another alternative, the feedstock is not heated by indirect heating of at least part of the effluent produced by stage c). In that case, the feedstock is mixed with that portion of the recovered effluent resulting from stage c), the portion of the recovered effluent resulting from stage c) having substantially the same temperature as the feedstock or Lower temperature than raw material.

The further heating stream advantageously consists of a hydrogen-rich gas effluent originating from the hydrogen supply and/or a gas effluent produced by separation stage c). The hydrogen-rich gas effluent originating from the hydrogen supply and/or at least a part of the gas effluent produced by separation stage c) is advantageously in a mixture with at least a part of the liquid effluent produced by stage c) Or separately injected upstream of stage a) and/or stage b). The hydrogen-rich gas stream may therefore advantageously be preheated in a mixture with at least a portion of the liquid effluent or separately before mixing, preferably by optionally passing through at least one exchanger and/or at least one oven or any other heating element known to those skilled in the art. Stage c2)

According to the invention, the treatment method comprises a separation stage c2), which is advantageously carried out in at least one washing/separation section fed with the first gas effluent and another part of the liquid effluent produced by stage c1) and an aqueous solution, this stage is carried out at a temperature of 20 to less than 200°C and at a pressure substantially equal to or less than the pressure of stage b), so as to obtain at least a gas effluent, a first aqueous effluent and a hydrocarbon effluent .

This separation stage c2) is carried out in at least one separation drum, called a high-pressure or medium-pressure cryogenic separation drum, also known to those skilled in the art by the name CHPS (cold high-pressure separator). Therefore, it is better to use a "cold high-pressure" separator in this stage c2), and the pressure is essentially equal to the operating pressure of stage b). The term "substantially equal to the pressure of stage b)" is to be understood as meaning a pressure of stage b) whose pressure difference relative to the pressure of stage b) is 0 to 1 MPa, preferably 0.005 to 0.3 MPa and particularly preferably 0.01 to 0.01 MPa. 0.2 MPa. Preferably, the pressure in stage c2) is the pressure reduced due to the pressure drop in stage b). The fact that at least part of the separation stage c2) is operated at a pressure substantially equal to the operating pressure of stage b) further facilitates the recovery of hydrogen.

The separation stage c2) can also be carried out at a lower pressure than in stage b).

The separation stage c2) may also comprise a (first) separation stage at a pressure substantially equal to the operating pressure of stage b), followed by a temperature equal to or lower than each previous separation stage of stage c2) and at least one other separation stage performed at a lower pressure.

The temperature at which the separation in stage c2) is carried out is between 20 and less than 200°C, preferably between 25 and 120°C and particularly preferably between 30 and 70°C.

It is important to operate within this temperature range (and thus not overcool the hydrotreated effluent) at risk of blockage in the line due to precipitation of ammonium chloride salts.

The washing/separating section of stage c2) may be made at least partly from common or separate items of washing and separation equipment, which are well known (separating drums, pumps, heat exchangers operating at various pressures and temperatures vessels, scrubbers and the like). This separation stage c2) can be carried out in particular in the same manner as stage c) described above.

When one (or two) hydrocracking stages (described below) are present, this stage c2) may additionally be fed with at least a portion of the hydrocracked effluent produced by the optional hydrocracking stage g).

At least a portion of the hydrocarbon effluent produced from stage c2) may be recovered as a liquid quench upstream of stage a) and/or stage b) and/or stage g). The injection of the hydrocarbon effluent produced from stage c2) can be in the first catalytic bed of the reaction section of stage a) and/or stage b) and/or stage g) or between the different catalytic beds of each section. conduct. When the hydrogenation reaction section of stage a) includes two reactors operating in a replaceable mode, at least a portion of the hydrocarbon effluent produced from stage c2) can be recovered between the two reactors. Stage d) of separating H ₂ S contained in the first aqueous effluent

According to the invention, the treatment method includes a stage d) of separating the H ₂ S contained in the first aqueous effluent, so as to obtain a gas phase containing H ₂ S and a second aqueous effluent, the gas containing H ₂ S The phase is preferably recovered at least partially upstream of stage a) and/or stage b) and/or stage g).

Stage d) of separating the _H2S contained in the first aqueous effluent is advantageously carried out by stripping in a stripping column using a gas stream, preferably an inert gas stream. A stripper is a distillation column in which a gas stream, preferably an inert gas stream, is injected at the bottom. A gaseous phase containing _H2S is recovered at the top of the column and a second aqueous effluent is recovered at the bottom of the column. The inert gas stream may be hydrogen, nitrogen or steam. Preferably, stage d) is carried out by steam stripping.

Stripping using a gas stream, preferably an inert gas stream, allows obtaining very low levels of dissolved _H2S in the second aqueous effluent at the bottom of the stripper.

The stage d) of separating H ₂ S is generally carried out at a pressure of 0.5 to 1.5 MPa, preferably 0.5 to 1 MPa and particularly preferably 0.6 to 0.9 MPa.

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

The flow rate of the inert gas stream is generally such that the flow rate of the inert gas stream (expressed in standard m ³ /hour (Sm ³ /h)) and the flow rate of the first aqueous effluent to be treated (under standard conditions (15°C, 0.1 MPa) The ratio (expressed in m ³ /hour below) is 50 to 600 Sm ³ /m ³ , preferably 200 to 400 Sm ³ /m ³ . The standard m ³ should be understood as meaning the amount of gas in a volume of 1 m ³ at 0°C and 0.1 MPa.

Stripping under these operating conditions makes it possible to specifically separate _H2S from the first aqueous effluent without entraining _NH3 (which remains mainly in the aqueous phase).

According to an alternative form, in stage d) a part of the gas phase at the top of the stripping column, which contains H ₂ S and the inert gas stream (steam), is condensed and preferably at least partially liquid refluxed and reinjected to the upper part of the stripping tower. Condensation is generally performed by cooling to a temperature of 30 to 65°C, for example using cold water.

This liquid reflux allows the temperature at the top of the stripper to be controlled/lowered.

The H ₂ S-containing gas phase withdrawn at the top of the stripper is at least partly upstream of stage a) and/or stage b) and/or also in the hydrocracking stage g) and/or in the selective hydrogenation of the pretreatment Stage a0) is recycled upstream (when present) in order to serve as vulcanizing agent for the catalyst. Before its recovery, it may undergo at least one further purification stage, such as contact with a liquid or washing with an amine.

According to another alternative, stage d) of separating the H ₂ S contained in the first aqueous effluent can also be carried out by liquid/liquid extraction, in which an inert solvent or reactant is brought into contact with the aqueous effluent. Stage e) of separating NH ₃ contained in the second aqueous effluent

According to the invention, the treatment method includes stage e) of separating NH ₃ contained in the second aqueous effluent produced by stage d), so as to obtain a gas phase containing NH ₃ and a third aqueous effluent containing The gas phase of NH ₃ is preferably recovered at least partially upstream of stage a) and/or stage b) and/or stage g).

Stage e) of separating the NH ₃ contained in the second aqueous effluent is advantageously carried out by stripping with an inert gas stream in a stripping column. A gaseous phase containing _NH3 is recovered at the top of the column and a third aqueous effluent is recovered at the bottom of the column. The inert gas stream may be hydrogen, nitrogen or steam. Preferably, stage e) is carried out by steam stripping.

Stripping with an inert gas stream makes it possible to obtain very low levels of dissolved _NH3 at the bottom of the stripper, allowing recovery of a third aqueous effluent that can be introduced into conventional wastewater treatment.

The stage e) of separating NH ₃ is generally carried out at a pressure of 0.1 to less than 0.5 MPa, preferably 0.05 to 0.2 MPa.

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

The flow rate of the inert gas stream is generally such that the flow rate of the inert gas stream (expressed in standard m ³ /hour (Sm ³ /h)) and the flow rate of the raw material to be treated (expressed in m ³ under standard conditions (15°C, 0.1 MPa) /hour) ratio is 50 to 600 Sm ³ /m ³ , preferably 200 to 400 Sm ³ /m ³ . The standard m ³ should be understood as meaning the amount of gas in a volume of 1 m ³ at 0°C and 0.1 MPa.

According to an alternative form, in stage e) a part of the gas phase at the top of the stripping column, which contains NH ₃ and the inert gas stream (steam), is condensed and preferably at least partially liquid refluxed and reinjected to the upper part of the stripping tower. Condensation is generally performed by cooling to a temperature of 30 to 65°C, for example using cold water.

This liquid reflux allows the temperature at the top of the stripper to be controlled/lowered.

According to an alternative form, the NH ₃ -containing gas phase can be at least partially upstream of stage a) and/or stage b) and/or stage g) and/or also upstream of selective hydrogenation stage a0) advantageously to suit The stoichiometric amount of salt formed during separation/washing stage c) is recovered. Fractionation stage f) ( optional )

The process according to the invention may comprise a stage of fractionating all or part (preferably all) of the hydrocarbon effluent produced in stage c) in order to obtain at least a gas stream and at least two liquid hydrocarbon streams, the two liquid hydrocarbon streams being At least a first hydrocarbon fraction containing compounds with a boiling point of less than or equal to 175°C (naphtha fraction), in particular 80 to 175°C, and a second hydrocarbon fraction (middle distillate) containing compounds with a boiling point greater than 175°C material fraction).

Stage f) makes it possible to specifically remove gases dissolved in the liquid hydrocarbon effluent, such as, for example, ammonia, hydrogen sulfide and light hydrocarbons having 1 to 4 carbon atoms.

This optional fractionation stage f) is advantageously carried out at a pressure less than or equal to 3.0 MPa abs., preferably from 0.5 to 2.5 MPa abs.

According to one embodiment, stage f) may be carried out in a section advantageously comprising at least one stripping column equipped with a reflux circuit including a reflux drum. The stripper is fed the liquid hydrocarbon effluent produced by stage c) and is fed a vapor stream. The liquid hydrocarbon effluent produced from stage c) is optionally heated before entering the stripper. Therefore, the lightest compounds are entrained overhead and into the reflux loop including the reflux drum where gas/liquid separation occurs. The gas phase containing light hydrocarbons is withdrawn from the reflux drum as a gas stream. Advantageously, a hydrocarbon fraction containing compounds with a boiling point less than or equal to 175° C. is withdrawn from the reflux drum. Advantageously, a hydrocarbon fraction containing compounds with a boiling point greater than 175° C. is withdrawn at the bottom of the stripping column.

According to other embodiments, this fractionation stage f) may employ a stripping column followed by a distillation column or only a distillation column.

Depending on the purpose or use of the fraction produced by fractionation stage f), one skilled in the art will adjust the cut point in the stripping and/or distillation operation. For example, the end point of the naphtha fraction may need to be adjusted to 150, 175 or 200°C.

The first hydrocarbon fraction comprising compounds having a boiling point less than or equal to 175°C and the second hydrocarbon fraction comprising compounds having a boiling point greater than 175°C, optionally mixed, may be wholly or partially sent to steam cracking Units, as a result of which olefins can be (re)formed to participate in the formation of polymers. Preferably, only a part of the fractions is sent to the steam cracking unit; at least a part of the remaining fraction is optionally recovered in at least one of the stages of the process and/or sent to a fuel storage unit, e.g. for A unit for storing naphtha, a unit for storing diesel oil, or a unit for storing kerosene, which is produced from conventional petroleum raw materials.

According to a preferred embodiment, the first hydrocarbon fraction containing compounds with a boiling point less than or equal to 175°C is completely or partially sent to a steam cracking unit, and the second hydrocarbon fraction containing compounds with a boiling point greater than 175°C is fed The hydrocarbon fraction is sent to the hydrocracking stage g) and/or to the fuel storage unit.

In a particular embodiment, this optional fractionation stage f) may make it possible to obtain, in addition to the gas stream, a naphtha fraction comprising compounds having a boiling point less than or equal to 175°C, preferably 80 to 175°C, comprising a boiling point greater than The middle distillate fraction contains compounds with boiling points of 175°C and less than 385°C and the hydrocarbon fraction containing compounds with boiling points greater than or equal to 385°C (called heavy hydrocarbon fractions). The naphtha fraction may be completely or partially sent to a steam cracking unit and/or a unit for storing naphtha produced from conventional petroleum feedstocks; it may also be recovered; the middle distillate may also be The fractions are fed completely or partially to steam cracking units or units for the storage of diesel fuel produced from conventional petroleum feedstocks or recovered; the heavy fractions (for their part) may be fed at least in their presence Partially sent to steam cracking unit or to hydrocracking stage g).

In another particular embodiment, this optional fractionation stage f) may make it possible to obtain, in addition to the gas stream, a naphtha fraction comprising compounds having a boiling point less than or equal to 175°C, preferably from 80 to 175°C, comprising Kerosene fractions containing compounds with boiling points greater than 175°C and less than or equal to 280°C, diesel fractions containing compounds with boiling points greater than 280°C and less than 385°C, and hydrocarbon fractions containing compounds with boiling points greater than or equal to 385°C ( called heavy hydrocarbon fraction). The naphtha fraction, kerosene fraction and/or diesel fraction may be sent completely or partially to a steam cracking unit or respectively to naphtha, kerosene or diesel pools produced from conventional petroleum feedstocks; or Recycle. This heavy fraction, as far as a part thereof is concerned, can be sent at least partially to a steam cracking unit or to a hydrocracking stage g) when present.

In another specific embodiment, the naphtha fraction produced from stage f) and comprising compounds having a boiling point less than or equal to 175°C is fractionated to obtain a heavy naphtha comprising compounds having a boiling point between 80 and 175°C. Fractions and light naphtha fractions containing compounds having a boiling point of less than 80° C., at least a part of the heavy naphtha fraction being fed to an aromatic complex, including at least one reforming naphtha for the purpose of preparing aromatic compounds Oil stage. According to this embodiment, at least a part of the light naphtha fraction is fed to the steam cracking stage h) described below.

The gas effluent produced by fractionation stage f) may constitute the subject of further purification and separation for the purpose of recovering at least light hydrocarbons, in particular ethane, propane and butane, which may be advantageously isolated or in a The mixture is fed to one or more furnaces of the steam cracking stage h) in order to increase the overall olefin yield. Hydrocracking stage g) ( optional )

According to an alternative form, the process of the invention may comprise a hydrocracking stage g) carried out after a separation stage c) using at least a part of the hydrocarbon effluent produced by stage c) or after using a gas containing a gas having a temperature greater than 175°C. A fractionation stage f) of at least a part of the second hydrocarbon fraction of boiling point compounds is carried out.

Advantageously, stage g) employs a hydrocracking reaction well known to those skilled in the art, and more particularly makes it possible to remove heavy compounds (for example having greater than A compound with a boiling point of 175°C) is converted into a compound with a boiling point less than or equal to 175°C. Other reactions may be carried out, such as hydrogenation of olefins or aromatics, hydrodemetallation, hydrodesulfurization, hydrodenitrification and the like.

These compounds with boiling points greater than 175°C have high BMCI and contain more naphthenic, naphthenic-aromatic and aromatic compounds relative to lighter compounds, thus resulting in higher C/H ratios. This high ratio is the cause of coking in the steam cracker, thus requiring a steam cracker specifically for this fraction. When it is desired to minimize the yield of these heavy compounds (middle distillate fraction) and maximize the yield of light compounds (naphtha fraction), these compounds can be at least partially converted into light compounds by hydrocracking. Compounds (generally beneficial to steam cracking unit fractions).

Therefore, the method of the present invention may comprise a hydrocracking stage g) carried out in a hydrocracking reaction section using at least one fixed bed reactor with n catalytic beds, n being an integer greater than or equal to 1, each Each catalytic bed contains at least one hydrocracking catalyst, and the hydrocracking reaction zone is fed with at least a portion of the hydrocarbon effluent produced by stage c) and/or is fed with a hydrocarbon effluent produced by stage f). At least a portion of the second hydrocarbon fraction of compounds having a boiling point greater than 175°C and a gas stream containing hydrogen, the hydrocracking reaction section is at an average temperature of 250 to 450°C and a hydrogen partial pressure of 1.5 to 20.0 MPa abs. and a space velocity per hour of 0.1 to 10.0 h ^-1 is used in order to obtain the first hydrocracked effluent.

Therefore, the hydrocracking reaction section is advantageously at an average temperature of 250 to 450°C, preferably 320 to 430°C, and a hydrogen partial pressure of 1.5 to 20.0 MPa abs., preferably 3 to 18.0 MPa abs. and It is adopted at hourly space velocity (HSV) of 0.1 to 10.0 h ^-1 , preferably 0.1 to 5.0 h ^-1 , preferably 0.2 to 4 h ^-1 . The hydrogen coverage in stage g) is advantageously between 80 and 2000 sm ^hydrogen / ^m fresh feed of stage a) and preferably between 200 and 1800 sm ^hydrogen /m ^fresh feed of stage a) between raw materials. The definitions of average temperature (WABT), HSV and hydrogen coverage correspond to those described above. Advantageously, the hydrocracking reaction zone is carried out at a pressure equivalent to the pressure used in the reaction zone of hydrogenation stage a) or hydrotreating stage b).

Advantageously, this stage g) is carried out in a hydrocracking reaction section comprising at least one, preferably one to five fixed bed reactors with n catalytic beds, n being greater than or equal to 1, preferably 1 to 10. In a preferred mode, an integer from 2 to 5, each of the beds includes at least one and preferably no more than ten hydrocracking catalysts. When the reactor includes several catalytic beds, that is at least two, preferably two to ten, and in a preferred way two to five catalytic beds, these catalytic beds are preferably arranged in series in the reaction in the vessel.

The hydrocracking effluent can be recovered at least partially in hydrogenation stage a) and/or in hydrotreating stage b) and/or in separation stage c). Preferably, it is recovered in separation stage c).

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

The second hydrocracking stage g') is carried out in the hydrocracking reaction section, which uses at least one fixed bed reactor with n catalytic beds, where n is an integer greater than or equal to 1, and each catalytic bed has Comprising at least one hydrocracking catalyst, the hydrocracking reaction zone is fed at least a part of the first hydrocracked effluent produced by the first hydrocracking stage g) and a gas stream comprising hydrogen, The hydrocracking reaction section is adopted at an average temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa abs., and an hourly space velocity of 0.1 to 10.0 h ^-1 in order to obtain the second hydrogenated Effluent from lysis. The preferred operating conditions and the catalysts for use 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 the same or different.

The second hydrocracking stage is preferably carried out in a hydrocracking reaction section comprising at least one, preferably one to five fixed bed reactors with n catalytic beds, n being greater than or equal to 1, preferably From 1 to 10, in a preferred mode an integer from 2 to 5, each of the beds includes at least one and preferably no more than ten hydrocracking catalysts.

The operating conditions used in the hydrocracking stage are generally such that per pass conversion to at least 80% by volume is achieved with a temperature of less than or equal to 175°C, preferably less than 160°C and in a preferred manner less than 150°C. The product of a compound with a boiling point of ℃ is greater than 15% by weight and 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 naphtha fraction (having a temperature of less than or equal to 175°C, especially 80 to less than or The selectivity of compounds equal to the boiling point of 175°C). This single-pass conversion is limited by the use of high recoveries in the loop for the second hydrocracking stage. The ratio is defined as the ratio of the feed flow rate of stage g') to the flow rate of the feed material of stage a) or stage b); preferably, the ratio is between 0.2 and 4, in a preferred way between 0.5 and between 2.5.

The hydrocracking effluent from the second hydrocracking stage g') can be recovered at least partially in the hydrogenation stage a) and/or in the hydrotreating stage b) and/or in the separation stage c). Preferably, it is recovered in separation stage c).

Therefore, this hydrocracking stage does not necessarily make it possible to convert all hydrocarbon compounds having a boiling point greater than 175°C (middle distillate fraction) into hydrocarbon compounds having a boiling point less than or equal to 175°C (naphtha fraction). Therefore, after fractionation stage f) there may still be present a more or less significant proportion of compounds having a boiling point greater than 175°C. In order to increase the conversion, at least a part of this unconverted fraction can be introduced into the second hydrocracking stage g'). The other part can be drained and removed. Depending on the operating conditions of the process, the vent may be from 0 to 10% by weight of the fraction containing compounds having a boiling point greater than 175 °C relative to the input feedstock, and preferably from 0.5 to 5% by weight.

According to the invention, the hydrocracking stage is operated in the presence of at least one hydrocracking catalyst.

The hydrocracking catalyst used in the hydrocracking stage is a conventional hydrocracking catalyst of the bifunctional type known to those skilled in the art, which combines an acid function with a hydrogenation-dehydrogenation function and optionally at least A bonding matrix. The acid functionality is provided by supports with high specific surface areas (typically 150 to 800 m ² /g) exhibiting surface acidity, such as combinations of halogenated (especially chlorinated or fluorinated) alumina, boron and aluminum oxides, non- Crystalline silica-alumina and zeolite) contribution. The hydrogenation-dehydrogenation function is contributed by at least one metal selected from group VIB of the periodic table and/or at least one metal selected from group VIII.

Preferably, the hydrocracking catalyst contains a hydrogenation-dehydrogenation function, which contains at least one metal from Group VIII, which is selected from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum, and is preferably selected from Cobalt and nickel. Preferably, the catalyst also contains at least one metal from group VIB selected from chromium, molybdenum and tungsten (alone or in mixture) and preferably selected from molybdenum and tungsten. NiMo, NiMoW or NiW type hydrogenation-dehydrogenation functionalities are preferred.

Preferably, the content of metals from Group VIII in the hydrocracking catalyst is advantageously between 0.5% and 15% by weight and preferably between 1% and 10% by weight, with these percentages expressed as oxides Expressed as a weight percentage (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 metals from Group VIB in the hydrocracking catalyst is advantageously between 5% and 35% by weight and preferably between 10% and 30% by weight, with these percentages expressed as oxides. Expressed as a weight percentage (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 may also optionally include at least one accelerator element deposited on the catalyst and selected from the group consisting of phosphorus, boron and silicon, optionally at least one element from Group VIIA (preferably chlorine and fluorine), optionally In the case of at least one element from group VIIB (preferably manganese) and optionally at least one element from group VB (preferably lithium).

Preferably, the hydrocracking catalyst contains at least one selected from the group consisting of alumina, silica, silica-alumina, aluminate, alumina-boria, magnesium oxide, silica-magnesium oxide, oxide Amorphous or poorly crystalline porous inorganic matrices of the oxide type of zirconium, titanium oxide or clay (alone or in mixtures) and preferably alumina or silica-alumina (alone or in mixtures).

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

Preferably, the hydrocracking catalyst also optionally includes a zeolite selected from Y zeolite, preferably USY zeolite, alone or with a zeolite selected from β, ZSM-12, IZM-2, ZSM-22, Other zeolite combinations of ZSM-23, SAPO-11, ZSM-48 or ZBM-30 zeolite (alone or in mixture). Preferably, the zeolite is USY zeolite alone.

In the case where the catalyst contains a zeolite, the content of zeolite in the hydrocracking catalyst is advantageously between 0.1% and 80% by weight, preferably between 3% and 70% by weight, these percentages Expressed as a percentage of zeolite relative to the total weight of the catalyst.

Preferred catalysts include at least one metal from Group VIB and optionally at least one non-noble metal from Group VIII, at least one accelerator element (and preferably phosphorus), at least one Y zeolite and at least one alumina binder And preferably consists of it.

Even better catalysts include and preferably consist of nickel, molybdenum, phosphorus, USY zeolite, and optionally 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.

The hydrocracking catalyst is, for example, in the form of an extrudate.

In an alternative form, the hydrocracking catalyst employed in the second hydrocracking stage comprises a hydrogenation-dehydrogenation functionality comprising at least one noble metal from group VIII selected from the group consisting of palladium and platinum, alone or in mixtures. The content of precious metals from group VIII is advantageously between 0.01% and 5% by weight and preferably between 0.05% and 3% by weight, these percentages being expressed as the total amount of oxide (PtO or PdO) relative to the catalyst. Expressed as a percentage of weight.

According to another aspect of the invention, the hydrocracking catalyst additionally contains one or more organic compounds containing oxygen and/or nitrogen and/or sulfur. This catalyst is often referred to by the term "additive catalyst". Generally, the organic compound is selected from the group consisting of one or more carboxylic acids, alcohols, thiols, thioethers, sulfates, sulfoxides, ethers, aldehydes, ketones, esters, carbonates, amines, nitriles, and imines. , oxime, urea and amide functional chemical functional compounds or compounds including furan rings or sugars.

The preparation of catalysts for the hydrogenation, hydrotreating and hydrocracking stages is known and generally involves impregnating the support with metals from group VIII and from group VIB, when present, and optionally phosphorus and/or The boron stage is followed by drying and then optional calcination. In the case of additive catalysts, the preparation is generally carried out by simple drying without calcination after the introduction of the organic compound. The term "calcination" is here understood to mean heat treatment in a gas containing air or oxygen at a temperature greater than or equal to 200°C. Before their use in this stage of the process, the catalysts are typically subjected to vulcanization in order to form active entities. The catalyst of stage a) may also be a catalyst used in its reduced form, thus involving a reduction stage in its preparation.

The hydrogen-containing gas stream feeding the selective hydrogenation, hydrogenation, hydrotreating and hydrocracking reaction section may consist of a supply of hydrogen and/or may consist of recovered hydrogen produced in particular by separation stage c). Preferably, a further gas stream comprising hydrogen is advantageously introduced at the inlet of each reactor (especially operated in series) and/or at the inlet of each catalytic bed starting from the second catalytic bed of the reaction section. . These additional gas streams are also called cooling streams. It allows control of the temperature in a reactor in which the reactions taking place are generally highly exothermic.

The hydrocarbon effluent or the hydrocarbon stream obtained by treating plastic pyrolysis oil according to the method of the invention thus exhibits a composition compatible with the specifications of the feedstock at the inlet of the steam cracking unit. In particular, the composition of the hydrocarbon effluent or the hydrocarbon stream is preferably such that: -The total content of metal 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 in a preferred manner small At or equal to 0.5 ppm by weight, where: The silicon (Si) element content is less than or equal to 5.0 ppm (by weight), preferably less than or equal to 1 ppm (by weight) and in a preferred way less than or equal to 0.6 ppm (by weight) plus/ or the elemental iron (Fe) content is less than or equal to 200 ppb (by weight), - a sulfur content of less than or equal to 100 ppm (by weight), preferably less than or equal to 50 ppm (by weight), and/or - a nitrogen content of less than or equal to 100 ppm (by weight), preferably less than or equal to 50 ppm (by weight) and in a preferred manner less than or equal to 5 ppm (by weight), and/or -Asphaltene content is less than or equal to 5.0 ppm by weight, and/or -The total content of chlorine is less than or equal to 10 ppm (by weight), preferably less than 1.0 ppm (by weight), and/or -The mercury content is less than or equal to 5 ppb by weight, preferably less than 3 ppb by weight, and/or -The content of olefin compounds (monoolefins and diolefins) is less than or equal to 5.0% by weight, preferably less than or equal to 2.0% by weight and in a preferred manner less than or equal to 0.1% by weight.

The contents are given as relative weight concentration, weight percent (%), parts per million by weight (ppm) or parts per billion by weight (ppb) relative to the total weight of the stream under consideration.

The method according to the invention thus makes it possible to treat plastic pyrolysis oil in order to obtain an effluent, which can be injected completely or partially into a steam cracking unit. Stage of adsorbing heavy metals ( optional )

Any gas effluent and/or any liquid effluent produced from at least one of the separation stages c), d) or e) or the fractionation stage f) may be subjected to an optional stage of adsorption of heavy metals.

Such gas effluent may in particular be the gas effluent produced by stage c) and/or the _H2S -containing gas phase produced by stage d) and/or the _NH3 -containing gas produced by stage e) phase and/or gas effluent resulting from fractionation stage f).

Such liquid effluent may in particular be a hydrocarbon effluent produced from stage c) and/or the first and/or second hydrocarbon fraction produced from stage f).

This optional adsorption stage makes it possible to eliminate or reduce the amount of metallic impurities that may be present in the gaseous and liquid effluents, in particular heavy metals such as arsenic, zinc, lead and especially mercury. Such metal impurities, especially heavy metals, are present in the raw materials. Some impurities (especially based on mercury) can be converted in one of the stages of the method according to the invention. Its transformed form is easier to capture. Their elimination or reduction may be particularly necessary when at least a portion of such gaseous and liquid effluents is intended to be sent directly or after having been subjected to one or more optional further stages (such as fractionation stage f))) to a facility with strict specifications for metallic impurities stages, such as the steam cracking stage.

Therefore, an optional stage of adsorption of gaseous effluents and/or hydrocarbon effluents produced by the process according to the invention is advantageous especially if at least one of these effluents or the feedstock respectively contains more than 20 ppb (to weight), in particular when there are more than 15 ppb (by weight) of metallic elements of heavy metals (As, Zn, Pb, Hg and the like) and especially when at least one of these effluents or the feedstock respectively Contains more than 10 ppb (by weight) of mercury, and more particularly more than 15 ppb (by weight) of mercury.

This optional adsorption stage is advantageously carried out at a temperature of 20 to 250°C, preferably 40 to 200°C, and a pressure of 0.15 to 10.0 MPa abs., preferably 0.2 to 1.0 MPa abs.

This optional adsorption stage can be carried out by any adsorbent known to those skilled in the art which makes it possible to reduce the amount of such contaminants.

According to an alternative form, the optional adsorption stage is carried out in an adsorption section operated in the presence of at least one adsorbent comprising a porous support and at least one element which may be based on elemental form or on a metal sulfide or metal. Reactive phase of sulfur in oxide form or in metallic form in elemental form.

The porous support may be selected indiscriminately from the family of alumina, silica-alumina, silica, zeolites and/or activated carbon. Advantageously, the porous support system 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, and more preferably between 150 and 320 m ² /g. The specific surface area of the adsorbent is the surface measured by the BET method as described above.

The active phase is based on sulfur in elemental form, or in the form of metal sulfides or metal oxides, or in the metallic form of elemental form. Preferably, the active phase is in the form of a metal sulfide, in particular a sulfide of a metal selected from the group of copper, molybdenum, tungsten, iron, nickel or cobalt.

Advantageously, the active phase of the adsorbent contains from 1 to 70% by weight, preferably from 2 to 25% and most preferably from 3 to 20% by weight of sulfur, relative to the total weight of the adsorbent.

Advantageously, the weight proportion of metal relative to the total weight of the adsorbent is generally 1% to 60%, preferably 2% to 40%, in a preferred manner 5% to 30%, preferably 5% to 20%.

The residence time in this adsorption zone is generally between 1 and 180 minutes.

The adsorption section may include one or more adsorption towers. When the adsorption section includes two adsorption towers, one mode of operation may be "swing" operation, where one of the towers is in-line, ie in operation, while the other tower is in reserve. Another mode of operation is to have at least two columns operated in series in a replaceable mode.

Preferably, the adsorption section includes an adsorption tower for gas effluent and an adsorption tower for liquid effluent. Steam cracking stage h) ( optional )

The hydrocarbon effluent produced by the separation stage c) or at least one of the two liquid hydrocarbon streams produced by the optional stage f) can be sent completely or partially to the steam cracking stage h).

Advantageously, the gas effluent produced from the separation stage c) and/or the fractionation stage f) and containing ethane, propane and butane can also be sent completely or partially to the steam cracking stage h).

This steam cracking stage h) is advantageously carried out in at least one pyrolysis furnace at a temperature of 700 to 900°C, preferably 750 to 850°C and a relative pressure of 0.05 to 0.3 MPa. The residence time of these hydrocarbon compounds is generally less than or equal to 1.0 seconds (expressed as s), preferably between 0.1 and 0.5 s. Advantageously, the steam is introduced upstream of the optional steam cracking stage h) and after the separation stage c) (or fractionation stage f)). At the entrance to stage h), the amount of water introduced, advantageously in the form of steam, is advantageously between 0.3 and 3.0 kg water/kg hydrocarbon compound. Preferably, this optional stage h) is carried out in parallel in several pyrolysis furnaces in order to adapt the operating conditions to the various streams feeding stage h), in particular resulting from stage f), and also to manage the decoking time of the pipelines . The furnace includes one or more tubes arranged in parallel. A furnace may also represent a group of furnaces operating in parallel. For example, the furnace may be dedicated to the cracking of hydrocarbon fractions containing compounds with boiling points less than or equal to 175°C.

Effluents from various steam cracking furnaces are generally recombined prior to separation for purposes of effluent composition. It will be understood that the steam cracking stage h) includes the steam cracking furnace but also sub-stages related to steam cracking which are well known to those skilled in the art. Such sub-stages may include inter alia heat exchangers, columns and catalytic reactors and recyclings to the furnace. The column generally makes it possible to fractionate the light fractions for the purpose of recovering at least a light fraction containing hydrogen and compounds having 2 to 5 carbon atoms, and a fraction containing pyrolysis oil, and optionally a fraction containing pyrolysis oil. effluent. The column makes it possible to separate the various components of the fractionated light fraction in order to recover at least an ethylene-rich fraction (C ₂ fraction) and a propylene-rich fraction (C ₃ fraction) and optionally a butene-rich fraction (C ₄ fraction) . These catalytic reactors enable the hydrogenation of C ₂ , C ₃ and indeed even C ₄ fractions and pyrolysis petroleum in particular. It is advantageous to recycle saturated compounds, especially those having 2 to 4 carbon atoms, to the steam cracking furnace in order to increase the overall yield of olefins.

This steam cracking stage h) enables satisfactory contents, in particular greater than or equal to 30% by weight, in particular greater than or equal to 40% by weight, relative to the weight of the steam cracking effluent under consideration, and indeed even greater than or equal to Equivalent to 50% by weight of total olefins containing 2, 3 and 4 carbon atoms to obtain at least one olefin containing 2, 3 and/or 4 carbon atoms (i.e. C ₂ , C ₃ and/or C ₄ olefins) effluent. These C ₂ , C ₃ and C ₄ olefins can then advantageously be used as polyolefin monomers.

According to a preferred embodiment of the invention, the method for treating raw materials containing plastic pyrolysis oil preferably consists of connecting together the following stages, preferably in the given order: - Hydrotreating stage b), separation/washing Stage c), stage d) for separation of H ₂ S with recovery of H ₂ S in stage b), and stage e) for separation of NH ₃ - hydrogenation stage a), hydrotreatment stage b), separation/washing stage c) , stage d) of separation of H ₂ S with recovery of H ₂ S in stages a) and/or b), and stage e) of separation of NH ₃ - hydrogenation stage a), hydrotreatment stage b), separation/washing stage c), stage d) of separating H _{2 S by recovering H 2 S} in stages a) and/or b), and stage e) of separating NH ₃ and fractionation stage f) - hydrogenation stage a), hydrotreating stage b), separation/washing stage c), stage d) of separating H ₂ S to recover H ₂ S in stages a) and/or b), and stage e) of separating NH ₃ and fractionation stage f) and adding A hydrocarbon fraction containing compounds with a boiling point greater than 175° C. is introduced in hydrocracking stage g) and the hydrocracked effluent is recovered in stage c).

Furthermore, all embodiments may include and preferably consist of a pretreatment stage a0).

Furthermore, all embodiments may comprise and preferably consist of a steam cracking stage h). Analytical methods used

Analytical methods and/or standards for determining the characteristics of various streams, particularly feedstocks and effluents to be treated, are known to those skilled in the art. They are specifically listed below by way of information. Other methods considered equivalent may also be used, in particular equivalent IP, EN or ISO methods: Table 1 describe method Density at 15°C ASTM D4052 Sulfur content ISO 20846 Nitrogen content ASTM D4629 Acid value ASTM D664 Bromine value ASTM D1159 Diolefin content value of maleic anhydride MAV method(1) The content of oxygen-containing molecules Combustion + Infrared Content of paraffins UOP990-11 Content of cycloalkanes and olefins UOP990-11 Aromatic hydrocarbon content UOP990-11 Halogen content ASTM D7359 Chloride content ASTM D7536 Metal content: ASTM D5185 P Fe Si Na B simulated distillation ASTM D2887 (1) MAV method, which is described in the literature: 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, Volume 62 (2007) , Issue 1, pages 57-68 with a list of figures

The details of the elements mentioned in Figures 1 and 2 allow a better understanding of the present invention, which is not limited to the specific embodiments shown in Figures 1 and 2. The various embodiments presented may be used alone or in combination with each other, without limitation to the combinations.

Figure 1 represents a diagram of a specific embodiment of the process according to the invention, which process comprises: - stage a) of the hydrogenation of the (optional) hydrocarbon feedstock 1 produced in a hydrogen-rich gas 2 and optionally in a stream 3 Produced by the pyrolysis of the plastic in the presence of the contributed amine and, optionally, the vulcanizing agent contributed by stream 4 (especially at the beginning of the cycle); - hydrotreatment stage b), which is fed into the hydrogenation stage a) The hydrocarbon effluent 5 produced (if present), and the stream 6 feeding the hydrogen-rich gas; - separation stage c), which is fed the effluent 7 produced by the hydrotreating stage b) and in the aqueous in the presence of the solution 10 in order to obtain at least a gas effluent 11, a first aqueous effluent 12 containing _H2S , HCl and _NH3 and a hydrocarbon effluent 13; - separation of the _H2S contained in the first aqueous effluent 12 stage d), preferably by stripping with a steam-containing stream 19, which makes it possible to obtain a H ₂ S-containing gas phase 20 and a second aqueous effluent 21, which H ₂ S-containing gas phase is at least partially Recycling takes place upstream of stage a) and/or stage b), preferably upstream of stage a) when present. Such recovery of the H ₂ S-containing phase 20 makes it possible to maintain the catalyst of stages a) and/or b) in the form of sulfide and thus reduce the contribution of the sulfurizing agent 4; - the separation is contained in the second aqueous effluent 21 Stage e) of NH ₃ is preferably carried out by stripping with a steam-containing stream 19 so that a gas phase 22 and a third aqueous effluent 23 containing NH ₃ can be obtained.

FIG. 2 shows a diagram of another specific embodiment of the method of the invention based on the diagram of FIG. 1 . The diagram includes stage c) which is carried out in two stages and is further followed by a fractionation stage f) and a hydrocracking stage g).

The hydrogenation stage a) and the hydrotreating stage b) are carried out as described in Figure 1 . This separation stage c), carried out in two stages, consists in particular: stage c1) of the separation of the hydrotreated effluent 7, carried out at high pressure and high temperature (HHPS), in order to obtain at least a gaseous effluent 8 and a liquid effluent 9, a part 9a of which can be recovered (not shown) upstream of stage a) or upstream of stage b), - separation stage c2), which is carried out at high pressure and low temperature (CHPS) and fed to the gas effluent 8 and the The further part 9b of the liquid effluent and the aqueous solution 10 produced by stage c1) makes it possible to obtain a gaseous effluent 11 containing at least hydrogen, an aqueous effluent 12 containing dissolved salts and dissolved _H2S and _NH3 , and hydrocarbon effluents13.

The stage d) of separating H ₂ S and the stage e) of separating NH ₃ are carried out as described in Figure 1 . The recovery of phase 20 containing _H2S is carried out in the same way. It can also be recovered at least partially in hydrocracking stage g).

Optionally, a stage f) of fractionation of the hydrocarbon effluent 13 is carried out, which makes it possible to obtain at least a gas effluent 14, a first hydrocarbon fraction 15 (naphtha fraction) comprising compounds with a boiling point less than or equal to 175° C. and The second hydrocarbon fraction 16 (middle distillate fraction) contains compounds having a boiling point greater than 175°C.

At the end of stage f), a part of this first hydrocarbon fraction 15 comprising compounds with a boiling point less than or equal to 175° C. can be sent to a steam cracking process (not shown). Another part of the first hydrocarbon fraction 15 can be fed to the hydrogenation stage a) and/or the hydrotreating stage b) (recovery not shown).

In Figure 2, at least a portion of the second hydrocarbon fraction 16 produced by stage f) and containing compounds having a boiling point greater than 175° C. is fed to a hydrocracking stage g), which is formed in at least one hydrocracking stage containing at least one The reaction is carried out in a fixed bed reactor with catalyst and hydrogen is fed 17. The hydrocracked effluent 18 can be recovered between the separation stages c1) and c2) or upstream of the separation stage c) (not shown).

Instead of injecting the amine stream 3 at the inlet of the hydrogenation stage a), it is possible to inject the amine stream 3 at the inlet of the hydrotreating stage b), the inlet of the separation stage c), the inlet of the hydrocracking stage g) (when these phase exists), or not, depending on the characteristics of the raw material.

Only the main stages with the main streams are shown in Figures 1 and 2 to enable a better understanding of the invention. It should be clearly understood that all items of equipment required for operation are present (drums, pumps, exchangers, ovens/furnaces, towers and the like), even if not shown. It will also be understood that the hydrogen-rich gas stream (supply or recycle) as described above may be injected at the inlet of each reactor or catalytic bed or between two reactors or two catalytic beds. Components for purifying and recovering hydrogen well known to those skilled in the art may also be used. Example Example 1 ( according to the invention )

The raw material 1 treated in this method at a flow rate of 10 000 kg/h (10 T/h) is a plastic pyrolysis oil (that is, containing 100% by weight of the plastic pyrolysis oil), which exhibits the performance indicated in Table 2 characteristics. Table 2 : Characteristics of raw materials describe / method unit Pyrolysis oil Density at 15°C ASTM D4052 g/cm ³ 0.845 Sulfur content ISO 20846 ppm by weight 170 Nitrogen content ASTM D4629 ppm by weight 2264 Acid value ASTM D664 mg KOH/g 3.95 Bromine value ASTM D1159 gBr ₂ /100 g 90 Diolefin content value of maleic anhydride MAV method ⁽¹⁾ % by weight 36 Elemental oxygen content ASTM D5622 % by weight 1.0 Content of paraffins UOP990-11 % by weight 19 Content of cycloalkanes UOP990-11 % by weight 7 Alkene content UOP990-11 % by weight 31 Aromatic hydrocarbon content UOP990-11 % by weight 43 Halogen content ASTM D7359 ppm by weight 400 Asphaltene content IFP 9313 ppm by weight 380 Metal content ASTM D5185 P ppm by weight 46 Fe ppm by weight 3 Si ppm by weight 130 Na ppm by weight 2.5 B ppm by weight 5 As ppb by weight 200 Hg ppb by weight 20 simulated distillation ASTM D2887 0% ℃ 45 10% ℃ 94 30% ℃ 135 50% ℃ 180 70% ℃ 250 90% ℃ 394 100% ℃ 550 (1) MAV method, which is described in the literature: 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, Volume 62 (2007) , Issue 1, pages 57 to 68

This feedstock 1 was subjected to a hydrogenation stage a) carried out in a fixed bed reactor in the presence of hydrogen 2 and a hydrogenation catalyst of the NiMo type on alumina under the conditions indicated in Table 3. Table 3 : Conditions for hydrogenation stage a) Reactor inlet temperature ℃ 290 Reactor outlet temperature ℃ 320 average temperature(WABT) ℃ 305 Hydrogen partial pressure MPa abs. 6.7 H ₂ /HC (hydrogen coverage in volume relative to the volume of feedstock) Sm ³ /m ³ 300 HSV (flow rate, based on volume of raw material/volume of catalyst) h ^-1 0.8

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

The degree of conversion observed (= (initial concentration - final concentration)/initial concentration) at the end of hydrogenation stage a) is indicated in Table 4. Table 4 : Conversion of entities during hydrogenation stage a ) Degree of conversion of dienes % ＞60 Degree of conversion of olefins % ＞60 Silicon retention % ＞75

The effluent 5 resulting from the hydrogenation stage a) is subjected directly without separation to the hydrotreatment stage b), which is present in a fixed bed in the presence of hydrogen and a hydrotreating catalyst of the NiMo type on aluminum oxide. Carry out under the conditions in Table 5. Table 5 : Conditions for hydrotreatment stage b) Average Hydrotreating Temperature (WABT) ℃ 325 Hydrogen partial pressure MPa abs. 6.5 H ₂ /HC (hydrogen coverage in volume relative to the volume of feedstock) Sm ³ /m ³ 400 HSV (flow rate, based on volume of raw material/volume of catalyst) h ^-1 0.5

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

The effluent 7 resulting from the hydrotreating stage b) is subjected to a separation stage c): a stream 10 of water is injected into the effluent resulting from the hydrotreating stage b); this mixture is subsequently used for washing the effluent containing Sulfur gas (sour gas) is processed in towers and separation drums to obtain gas fractions and liquid effluents. The yields of the various fractions obtained after the separation are shown in Table 6 (these yields correspond to the ratio of the amount by weight of the various products obtained relative to the weight of the feedstock upstream of stage a), expressed in percentage and Expressed as % w/w). Table 6 : Yield of various products obtained after isolation Gas fraction (NH ₃ + H ₂ S + H ₂ O + C ₁ -C ₄ ) %w/w 2.42 liquid fraction %w/w 99.31

All or part of the liquid fraction obtained can be subsequently upgraded in a steam cracking stage for the purpose of forming olefins, which can be polymerized for the purpose of forming recycled plastics.

This pyrolysis oil feedstock contains very small amounts of sulfur (170 ppm by weight). The sulfur present in the form of sulfur-containing molecules is hydrogenated in the reaction zone and converted into _H2S . This _H2S in the form of _H2S partial pressure ( _pH2Sp ) in the reactor helps maintain the sulfide phase of the alumina supported NiMo catalyst. However, the pH ₂ Sp obtained with this level of sulfur-containing compounds in the feedstock (170 ppm by weight) is insufficient to keep the catalysts in the sulfide phase throughout the cycle. If nothing is done, this causes the catalyst to rapidly deactivate. Therefore, it is recommended to add H ₂ S to the reaction system in order to achieve a sufficient pH ₂ Sp. This addition of _H2S may be in the form of injecting dimethyl disulfide (DMDS) into the pyrolysis oil feed at the inlet of the unit. DMDS readily decomposes to give CH ₄ and H ₂ S once it contacts the catalyst, thus generating sufficient pH ₂ Sp to keep the catalyst in the sulfide form. This mode of operation results in a high consumption of DMDS which is detrimental to the economics of the process.

Another way that forms the subject of the present invention is to use double stripping of the aqueous effluent to recover the H ₂ S discharged in the aqueous effluent and reinject this H ₂ S by dissolving in the pyrolysis oil feedstock. Entrance to the unit.

Injection of DMDS and/or recovery of _H2S at the inlet of the unit can be used to maintain adequate _pH2Sp of the reaction system but can also be used to neutralize any _NH3 produced by the hydrogenation of nitrogen-containing molecules. This is because the _H2S reacts with _NH3 to form ammonium sulfide, which will be almost completely washed out and transferred to the aqueous effluent (stream 12), allowing the gas stream at the top outlet of the stabilization column to be free of ammonia (stream 12). 11). Therefore, the gas stream without the presence of ammonia can be sent directly to the steam cracker in order to maximize olefin production.

The advantage of recycling the _H2S stream compared to injecting DMDS, both in the context of maintaining adequate _pH2Sp and in delivering a gas stream that has its ammonia removed, thus saving the amount of DMDS in the overall cycle.

The four operating conditions are shown in Table 7.

Case 1 : Simply strip the sulfurous water and inject only DMDS to maintain sufficient pH ₂ Sp to keep the catalyst in the sulfide phase.

Case 2 : Double-layer stripping of sulfur-containing water to recover the H ₂ S-dominated stream produced at the top of the first stripper at the inlet of the unit only to maintain sufficient pH ₂ Sp to keep the catalyst in the sulfurization In the physical phase. This situation is consistent with the present invention.

Case 3 : The sulfurous water is simply stripped and DMDS is injected in order to maintain sufficient pH ₂ Sp to keep the catalyst in the sulfide phase and also to neutralize all NH ₃ and deliver an NH ₃ -free gas stream.

Case 4 : Double stripping of sulfuric water to recover the H ₂ S dominant stream produced at the top of the first stripper at the inlet of the unit in order to maintain sufficient pH ₂ Sp to keep the catalyst in the sulfide phase Neutralize and also to neutralize all NH ₃ and deliver NH ₃ free gas stream. This situation is consistent with the present invention.

It can be observed that the present invention allows a saving of 19 kg/h DMDS when maintaining a minimum pH ₂ Sp is important to keep the catalyst in the sulfide form.

It can also be observed that the present invention enables even greater savings, namely a saving of 65 kg/h (75 to 10 = 65 kg/h) of DMDS for delivery of the ammonia-free gas phase. Table 7 : Operation conditions Case 1 Case 2 Case 3 Case 4 Simple steaming of sulfur-containing water to a minimum pH of ₂ Sp Double steaming of sulfuric water to a minimum pH of ₂ Sp Simple steaming and NH ₃ -free gas logistics of sulfur-containing water Double steaming and _NH3 -free gas flow of sulfur-containing water pH ₂ Sp at the outlet of the HDT reactor bar a 0.02 (0.002 MPa) 0.02 (0.002 MPa) 0.21 (0.021 MPa) 0.21 (0.021 MPa) DMDS injection flow rate kg/h 19 0 75 10 DMDS savings kg/h - 19 - 65 Operating conditions of stabilization tower Tower top pressure barg 7.1 (0.71 MPa) Tower bottom pressure barg 7.3 (0.73 MPa) Raw material inlet temperature °C 160 Number of theoretical boards 10 Heat flux of reboiler Gcal/h 0.55 Operating conditions of H ₂ S stripper Tower top pressure barg N/A 9 (0.9 MPa) N/A 9 (0.9 MPa) Tower bottom pressure barg 9.1 (0.91 MPa) 9.1 (0.91 MPa) Raw material inlet temperature °C 135 130 Number of theoretical boards - 12 12 Stripping steam flow rate kg/h 285 303 Operating conditions of NH ₃ stripping tower Tower top pressure barg N/A 2.4 (0.24 MPa) N/A 2.4 (0.24 MPa) Tower bottom pressure barg 2.6 (0.26 MPa) 2.6 (0.26 MPa) Raw material inlet temperature °C 124 126 Number of theoretical boards - 15 15 Stripping steam flow rate kg/h 535 538 Operating conditions of simple stripping tower (H ₂ S and NH ₃ ) Tower top pressure barg 2.4 (0.24 MPa) N/A 2.4 (0.24 MPa) N/A Tower bottom pressure barg 2.5 (0.25 MPa) 2.5 (0.25 MPa) Raw material inlet temperature ℃ 94 93 Number of theoretical boards - 15 15 Stripping steam flow rate kg/h 435 450 Stabilize the flow rate of gas at the top of the tower kg/h 177 169 179 153 The composition of the gas at the top of the stabilization tower H ₂ Mol% 29.3 31.7 26.4 33.7 H ₂ S Mol% without 0.2 1.8 2.1 NH ₃ Mol% 9.6 9.8 without without H ₂ O Mol% 1.4 1.4 1.4 1.4 C ₁ Mol% 33.8 29.9 46.7 35.6 C ₂ Mol% 6.9 7.3 6.8 7.9 C ₃ Mol% 3.9 4.0 3.7 4.4 C ₄ Mol% 10.7 11.1 10.1 11.9 C ₅₊ Mol% 4.5 4.4 3.1 3.1 Flow rate of H ₂ S-rich gas at the top of H ₂ S stripper kg/h N/A 14 N/A 44 Composition of the H ₂ S-rich gas from the top of the H ₂ S stripper H ₂ S Mol% N/A 92.2 N/A 92.2 NH ₃ Mol% 0.05 0.07 H ₂ O Mol% 7.74 7.72 Flow rate of sulfur-containing gas at the top of the sulfur-containing water stripping tower kg/h 40 twenty two 103 44 The composition of the acid gas at the top of the NH ₃ stripping (or simple stripping) tower H ₂ S Mol% 24.5 2.7 32.7 3.0 NH ₃ Mol% 44.3 66.1 35.7 65.7 H ₂ O Mol% 31.2 31.3 31.6 31.2 Stripping water flow rate kg/h 2553 2996 2551 3031 Composition of stripping water H ₂ O % by weight 99.9 99.9 99.9 99.9 NH ₃ ppm by weight ＜30 ＜30 ＜30 ＜30 H ₂ S ppm by weight <5 <5 <5 <5 hydrocarbon ppm by weight Trace amount Trace amount Trace amount Trace amount

1: Hydrocarbon feed 2: Stream of hydrogen-rich gas 3: Stream/amine stream 4: Stream 5: Effluent 6: Stream of hydrogen-rich gas 7: Effluent 8: Gas effluent 9: Liquid effluent 9a : One part of the liquid effluent 9b: Another part of the liquid effluent 10: Aqueous solution 11: Gas effluent 12: Aqueous effluent containing dissolved salts and dissolved H ₂ S and NH ₃ 13: Hydrocarbon effluent 14: Gas Effluent 15: first hydrocarbon fraction 16: second hydrocarbon fraction 17: hydrogen 18: hydrocracked effluent 19: stream containing steam 20: gas phase containing H ₂ S 21: second aqueous effluent 22: Gas phase 23 containing NH ₃ : third aqueous effluent a): stage b): stage c): stage c1): stage c2): stage d): stage e): stage g): stage f): stage

Figure 1 represents a diagram of a specific embodiment of the method of the present invention.

FIG. 2 shows a diagram of another specific embodiment of the method of the invention based on the diagram of FIG. 1 .

1: Hydrocarbon raw materials

2: Logistics of hydrogen-rich gas

3:Logistics/Amine Logistics

4:Logistics

5: Effluent

6: Logistics of hydrogen-rich gas

7: Effluent

10:Aqueous solution

11: Gas effluent

12: Aqueous effluent containing dissolved salts and dissolved H ₂ S and NH ₃

13: Hydrocarbon effluent

19: Logistics containing steam

20: Gas phase containing H ₂ S

21: Second aqueous effluent

22: Gas phase containing NH ₃

23:Third aqueous effluent

a): stage

b): stage

c): stage

d): stage

e): stage

Claims (18)

A method for processing a feedstock containing plastic pyrolysis oil, comprising: a) a hydrogenation stage optionally carried out in a hydrogenation reaction section using at least one fixed bed reactor with n catalytic beds, n being An integer greater than or equal to 1. Each catalytic bed contains at least one hydrogenation catalyst. The hydrogenation reaction section is fed with at least the raw material and a gas stream containing hydrogen. The hydrogenation reaction section is at an average temperature of 140 to 400°C. temperature, a hydrogen partial pressure of 1.0 to 10.0 MPa abs. and a space velocity per hour of 0.1 to 10.0 h ^-1 to obtain a hydrogenation effluent, b) in a hydrotreating reaction zone containing at least one hydrotreating catalyst The hydrotreating stage carried out in the hydrotreating reaction zone is fed with at least the feedstock or the hydrogenation effluent produced from stage a) and the gas stream containing hydrogen. The hydrotreating reaction zone is in An average temperature of 250 to 430°C, a hydrogen partial pressure of 1.0 to 10.0 MPa abs. and a space velocity per hour of 0.1 to 10.0 h ^-1 are used in order to obtain the hydrotreated effluent, c) separation stage, which is Feeding the hydrotreated effluent produced from stage b) and optionally the hydrocracked effluent produced from stage g) and the aqueous solution, in order to obtain at least a gas effluent, a first an aqueous effluent and a hydrocarbon effluent, d) the stage of separating the H ₂ S contained in the first aqueous effluent to obtain a gas phase containing H ₂ S and a second aqueous effluent, the gas containing H ₂ S the phase is recovered at least partially upstream of stage a) and/or stage b) and/or stage g) as the case may be, e) a stage of separation of NH ₃ contained in the second aqueous effluent in order to obtain a stage containing NH ₃ The gas phase and the third aqueous effluent, the gas phase containing NH ₃ is recovered at least partially upstream of stage a) and/or stage b) and/or stage g), as appropriate, f) optionally fractionated by Stage c) the stage of all or part of the hydrocarbon effluent produced in order to obtain at least a gas effluent and at least a first hydrocarbon fraction containing compounds with a boiling point less than or equal to 175°C and containing compounds with a boiling point greater than 175°C a second hydrocarbon fraction of the compound, g) optionally a hydrocracking stage carried out in a hydrocracking reaction zone using at least one fixed bed reactor with n catalytic beds, n being an integer greater than or equal to 1 , each catalytic bed comprising at least one hydrocracking catalyst, the hydrocracking reaction zone being fed with at least a portion of the hydrocarbon effluent produced by stage c) and/or fed with the hydrocarbon effluent produced by stage f) At least a portion of the second hydrocarbon fraction containing compounds having a boiling point greater than 175°C and a gas stream containing hydrogen, the hydrocracking reaction section is at an average temperature of 250 to 450°C and 1.5 to 20.0 MPa abs. A hydrogen partial pressure and a space velocity per hour of 0.1 to 10.0 h ^-1 are used in order to obtain the first hydrocracked effluent. The method of the preceding claim, wherein the H ₂ S-containing gas phase produced by stage d) is at least partially recovered upstream of stage a) and/or stage b) and/or stage g). A method as claimed in any one of the preceding claims, comprising the hydrogenation stage a). A method according to any one of the preceding claims, comprising the fractionation stage f). A method as claimed in any one of the preceding claims, comprising the hydrocracking stage g). A method according to any one of the preceding claims, wherein stage d) of isolating H ₂ S contained in the first aqueous effluent is by using a steam-containing stream at a pressure of 0.5 to 1 MPa and 80 to 150°C This is done by stripping the effluent at a temperature. A method according to any one of the preceding claims, wherein stage e) of separating NH ₃ contained in the second aqueous effluent is by using a steam-containing stream at a pressure of 0.1 to 0.5 MPa and a temperature of 80 to 150°C. This is done by stripping the effluent at temperature. The method according to any one of the preceding claims, wherein the separation stage c) includes the following stages: c1) A separation stage fed to the hydrotreated effluent produced by stage b), which stage is carried out at a temperature between 200 and 450°C and at a pressure substantially equal to the pressure of stage b) , in order to obtain at least gaseous effluent and liquid effluent, part of which is recycled upstream of stage a) and/or stage b), as appropriate, c2) A separation stage, which is fed the gaseous effluent produced by stage c1) and another part of the liquid effluent produced by stage c1) and an aqueous solution, which stage is at a temperature of 20 to less than 200°C temperature and at a pressure substantially equal to or less than the pressure of stage b), so as to obtain at least a gas effluent, a first aqueous effluent and a hydrocarbon effluent. A method as claimed in any one of the preceding claims, comprising pre-treating at least one stage a0) of a feedstock comprising plastic pyrolysis oil optionally in a mixture with the hydrocarbon effluent produced from stage c), said pre-treating The stage is carried out upstream of stage a) and/or stage b) and includes a filtration stage and/or a centrifugation stage and/or an electrostatic separation stage and/or a washing stage with the aid of an aqueous solution and/or an adsorption stage and/or selection. sexual hydrogenation stage. A method as claimed in any one of the preceding claims, wherein the hydrocarbon effluent produced by separation stage c) or at least one of the two liquid hydrocarbon fractions produced by stage f) is completely or partially sent to Steam cracking stage h) is carried out in at least one pyrolysis furnace at a temperature of 700 to 900° C. and a relative pressure of 0.05 to 0.3 MPa. A method as claimed in any one of the preceding claims, wherein the NH ₃ -containing gas phase produced by stage e) is at least partially recovered upstream of stage a) and/or stage b) and/or stage g). A method according to any one of the preceding claims, wherein a stream containing nitrogen compounds and/or sulfur compounds is injected upstream of stage a) and/or upstream of stage b). The method of any one of the preceding claims, wherein the hydrogenation catalyst includes a support selected from the group consisting of alumina, silica, silica-alumina, magnesium oxide, clay and mixtures thereof and includes at least one component from Group VIII The hydrogenation-dehydrogenation function of an element and at least one element from group VIB, or at least one element from group VIII. The method of any one of the preceding claims, wherein the hydrotreating catalyst includes a support selected from the group consisting of alumina, silica, silica-alumina, magnesium oxide, clay and mixtures thereof and includes Hydrogenation-dehydrogenation functionality of at least one element from group VIII and/or at least one element from group VIB. A method according to any one of the preceding claims, further comprising a second hydrocracking stage g') carried out in a hydrocracking reaction section, using at least one fixed bed reactor with n catalytic beds, n is an integer greater than or equal to 1, each catalytic bed comprising at least one hydrocracking catalyst, the hydrocracking reaction zone being fed with at least the first hydrocracking catalyst produced by the first hydrocracking stage g) A portion of the effluent of hydrogen cracking and a gas stream containing hydrogen. The hydrocracking reaction section is at a temperature of 250 to 450°C, a hydrogen partial pressure of 1.5 to 20.0 MPa abs., and a pressure of 0.1 to 10.0 h ^-1 per Hour space velocities are used in order to obtain a second hydrocracked effluent. The method of any one of the preceding claims, wherein the hydrocracking catalyst includes a support selected from the group consisting of halogenated alumina, a combination of boron and aluminum oxide, amorphous silica-alumina and zeolite and includes at least Hydrogenation-dehydrogenation functionality of a Group VIB metal selected from the group consisting of chromium, molybdenum and tungsten (alone or in mixtures) and/or at least one Group VIII metal selected from the group consisting of iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum . A product obtained by a method according to any one of claims 1 to 16. For example, the product of claim 17 includes: relative to the total weight of the product: Metallic elements in a total content less than or equal to 10.0 ppm (by weight), Contains iron at levels less than or equal to 200 ppb by weight, and/or Silicon in an amount less than or equal to 5.0 ppm (by weight), and/or Sulfur content less than or equal to 100 ppm by weight, and/or Nitrogen content less than or equal to 100 ppm (by weight), and/or Elemental chlorine in a content less than or equal to 10 ppm by weight, and/or Mercury content less than or equal to 5 ppb by weight.