Method for producing renewable kerosene by two-step hydroprocessing using a specific catalyst in the hydroconversion step with recycle of a heavy fraction

Abstract

The present invention describes a method for treating a feedstock originating from a renewable source comprising a hydrotreatment step a), a step b) of separation of at least one portion of the effluent resulting from step a) into at least one light fraction and at least one hydrocarbon effluent, a step c) of hydroconversion of the hydrocarbon effluent resulting from step c), using at least one specific bifunctional catalyst comprising at least one noble metal from group VIII selected from among platinum and palladium, and a support comprising at least one zeolite selected from among the MTW zeolites and the IZM-2 zeolite, alone or in a mixture, and at least one binder, a step d) of separation of the effluent resulting from step c), which makes it possible to separate at least one gaseous fraction and at least one hydrocarbon liquid effluent, and a step e) of fractionation of the hydrocarbon effluent resulting from step d) into at least one kerosene cut and at least one heavy cut having an initial boiling point of between 250 and 300°C, wherein all or part of the heavy cut resulting from the step e) is recycled to the hydroconversion step c).

Description

[0001] Description

[0002] Title: PROCESS FOR THE PRODUCTION OF RENEWABLE KEROSENE BY HYDROPROCESSING IN 2 STAGES USING A SPECIFIC CATALYST IN THE HYDROCONVERSION STAGE WITH RECYCLING OF A HEAVY FRACTION

[0003] Field of invention

[0004] The search for new renewable energy sources for fuel production is a major challenge in meeting both fuel demand and environmental concerns.

[0005] In this respect, the recovery of feedstocks from renewable sources into fuels has seen a very strong resurgence of interest in recent years. Examples of these feedstocks include vegetable oils (e.g. palm, rapeseed, soybean), animal fats, used cooking oils, oils of microbial origin (e.g. from algae), fish oils, long paraffins (waxes) from the Fischer-Tropsch process, whether crude or pre-treated, as well as mixtures of such feedstocks. These feedstocks mostly contain chemical structures such as triglycerides or esters or fatty acids, the structure and length of the hydrocarbon chain of the latter being compatible with the hydrocarbons present in diesel and kerosene.

[0006] One possible route is the catalytic transformation of the feedstock from renewable sources into deoxygenated paraffinic fuel in the presence of hydrogen (hydrotreatment). Many metal or sulfide catalysts are known to be active for this type of reaction.

[0007] These processes for hydrotreating feedstock from renewable sources are already well known and are described in numerous patents. Examples include patents: US 4,992,605, US 5,705,722, EP 1,681,337 and EP 1,741,768.

[0008] The use of solids based on transition metal sulfides allows the production of paraffins from ester-type molecules using two reaction pathways:

[0009] - hydrodeoxygenation leading to the formation of water by consumption of hydrogen and to the formation of hydrocarbons with carbon number (C n) equal to that of the initial fatty acid chains, - decarboxylation / decarbonylation leading to the formation of carbon oxides (carbon monoxide and dioxide: CO and CO2) and to the formation of hydrocarbons with one less carbon (C n -i) compared to the initial fatty acid chains.

[0010] The liquid effluent from these hydrotreatment processes, after separation, consists essentially of n-paraffins and is substantially free of sulfur, nitrogen and oxygen impurities. After hydrotreatment and gas separation, the sulfur content is typically between 1 and 20 ppm by weight, the nitrogen content is generally between 0.2 and 30 ppm by weight and the oxygen content is generally less than 2000 ppm by weight. Paraffins have a carbon atom number typically between 9 and 25, which is mainly dependent on the composition of the feedstock to be hydrotreated.

[0011] However, this liquid effluent cannot generally be incorporated as is into the kerosene or diesel pool, particularly due to insufficient cold properties and / or boiling temperatures that are too high. Indeed, the paraffins present lead to high pour points and therefore to freezing phenomena for uses at low temperatures. For example, eicosane (linear paraffin with 20 carbon atoms, C20H42) has a boiling point of 340°C and a melting point of 37°C. The boiling point of eicosane is therefore compatible with incorporation into a diesel pool, but its melting temperature can generate freezing problems and limit its use. As an illustration, the filterability limit temperature for winter diesel is a maximum of -15°C.Furthermore, the boiling temperature of eicosane makes it unincorporable into the kerosene pool, for which the final temperature of the distillation curve must be less than 300°C.

[0012] Depending on the incorporation rate and the preferred fuel pool (diesel or kerosene) being targeted, it may be necessary to carry out a hydroconversion step (hydroisomerization and / or hydrocracking reactions) to transform the linear paraffins in the hydrotreated liquid effluent. Hydroisomerization converts a linear paraffin into a branched paraffin while preserving the number of carbon atoms in the molecule. This improves the cold properties of the effluent because branched paraffins have better cold properties than linear paraffins. For example, nonadecane has a melting point of 32°C, while one of its monobranched isomers, 7-methyl-octadecane, has a melting point of -16°C. Hydrocracking converts a linear paraffin into linear or branched paraffins of lower molecular weight.This allows the effluent distillation curve to be adjusted as needed to make it compatible with the kerosene pool. For example, hydrocracking one eicosane molecule can produce two 2-methylnonane molecules. The boiling point of 2-methylnonane is 167°C, which is compatible with incorporation into the kerosene pool. The hydroconversion step is carried out on a bifunctional catalyst with both a hydro / dehydrogenating function and a Bronsted acid function. The operating conditions can be adapted to promote hydroisomerization or hydrocracking reactions as needed. In all cases, it is desirable to minimize the production of cracking products that are too light to be incorporated into the kerosene or diesel pool.

[0013] The appropriate choice of the acid phase promotes the isomerization of long linear paraffins and minimizes cracking. Thus, the shape selectivity of one-dimensional medium-pore zeolites (10 MR) such as ZSM-22, ZSM-23, NU-10, ZSM-48, ZBM-30 zeolites makes their use particularly suitable for obtaining isomerization-selective catalysts. Other acidic phases of zeolitic or non-zeolitic type such as halogenated aluminas (chlorinated or fluorinated in particular), phosphorus-containing aluminas, silica-aluminas or silica-containing aluminas can also be used.

[0014] However, it is well known that factors other than the acid phase impact the activity and selectivity of a bifunctional catalyst. The hydroisomerization and hydrocracking of normal paraffins have been the subject of numerous academic studies since the original work of Weisz or Coonradt and Garwood in the 1960s. The most commonly accepted mechanism involves first dehydrogenating the n-paraffin to an n-olefin in the hydrodehydrogenating phase and then, after diffusion to the acid phase, protonating it to a carbenium ion. After structural rearrangement and / or p-scission, the carbenium ions desorb from the acid phase as olefins after deprotonation. Then, after diffusion to the hydrodehydrogenating phase, the olefins are hydrogenated to form the final reaction products.It is then necessary to have a sufficiently active hydro / dehydrogenating function with respect to the acid function to, on the one hand, quickly supply the acid phase with olefins and, on the other hand, to quickly hydrogenate the olefinic intermediates after their reaction on the acid phase. This makes it possible, on the one hand, to maximize the activity of the catalyst and, on the other hand, to favor hydroisomerization compared to hydrocracking when the first reaction is desired, or to limit the production of cracking products that are too light when the hydrocracking reaction is desired. The use of a sufficiently active hydrogenating function is also desirable in order to limit the deactivation of the bifunctional catalyst by coking during the hydroconversion of n-paraffins (Alvarez et al., Journal of Catalysis, 162, 2, 179-189) for a range of fixed operating conditions.

[0015] The proximity between the two catalyst functions can also have an impact on the performance of the bifunctional catalyst. Thus, Zecevic et al. (Nature, 2015, 528, 245-254) recently studied the impact of platinum localization on the hydroisomerization performance of long paraffins (n-decane, n-nonadecane, pristane) of a bifunctional catalyst using USY zeolite as the acid phase and an alumina matrix. It is observed that the bifunctional catalyst for which platinum is deposited on alumina is systematically more selective in isomerization than the catalyst for which platinum is deposited in the zeolite. In view of these results, the skilled person is therefore inclined to favor localization of the hydrogenating function on the alumina matrix rather than on the acid phase to improve isomerization selectivity.From an activity point of view, the location of platinum on the alumina matrix has a variable impact depending on the long paraffin considered: positive impact with regard to n-decane, marginal impact with regard to n-nonadecane and finally negative impact with regard to pristane.

[0016] Noble metals (Pt, Pd) or group VI B transition metals (Mo, W) combined with group VIII transition metals (Ni, Co) can act as the hydrogenating function for the catalyst. Noble metals are used in their reduced form while transition metals are used in a sulfurized form. For the latter, there is a known synergistic effect between group VI B transition metals and group VIII transition metals, generally attributed to the decoration of group VI B sulfide phases by group VIII transition metals. These are referred to as molybdenum or tungsten sulfide phases promoted by nickel or cobalt ("CoMoS", "NiMoS", "NiWS"). This synergistic effect results in an increase in the catalytic activity of the promoted phase compared to a non-promoted phase.

[0017] The choice of the nature of the hydrogenating function, noble metal or sulfide type, depends on different criteria, of an economic nature (the price of noble metals is significantly higher than that of transition metals from groups VIA and VIII) or of a chemical nature (impact of the presence of contaminants). Thus, the hydrogenating activity of noble metals is higher than that of transition metal sulfides when the partial pressure of hydrogen sulfide (H2S) in the reaction medium is low or even zero. Conversely, the hydrogenating activity of transition metal sulfides is higher than that of noble metals when the partial pressure of H2S in the reaction medium becomes high (C. Marcilly, Acid-Base Catalysis, volume 2, 2003, Technip editions).

[0018] Patent US2022 / 0127537 teaches a process for hydrotreating a renewable feedstock. Said process comprises a step of hydrotreating the feedstock in the presence of hydrogen and a hydrotreating catalyst to deoxygenate said feedstock and thus produce a hydrotreated effluent. Said process comprises a step of hydroisomerization, in the presence of hydrogen and a hydroisomerization catalyst, of an effluent from the hydrotreated effluent to obtain a hydroisomerized effluent.The hydroisomerization catalyst used may comprise a group VIII metal selected from Pt and Pd, alone or in combination and a support which may be amorphous or crystalline selected from alumina, amorphous silica alumina, ferrierite, ALPO-31, SAPO-11, SAPO-31, SAPO, 37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-11, MeAPO-31, MeAPO-41, MgAPSO-11, MgAPSO-31, MgAPSO-41, MgAPSO-46, ELAPSO-11, ELAPSO-3, EMAPSO-41, ELAPSO-11, ELAPSO-31, ELAPSO-41, alone or in combination.

[0019] In a first embodiment claimed in claim 1, said process comprises a step of separating the hydrotreated effluent and / or the isomerized effluent to obtain a vapor effluent and a liquid effluent; a step of distilling the liquid effluent or the hydroisomerized effluent to produce a kerosene cut and a diesel cut and a step of hydrocracking the diesel cut to obtain a hydrocracked effluent comprising a kerosene cut. Unlike the present invention, said process therefore allows the coproduction of diesel and kerosene cuts which are separated by a distillation step and the diesel cut obtained then undergoes a hydrocracking step.

[0020] In a second embodiment claimed in claim 16, said process comprises obtaining a diesel cut from the hydroisomerized effluent, said diesel cut is characterized in that the normal paraffin concentration of a given carbon number is at least twice the corresponding normal paraffin concentration in said hydroisomerized effluent. Said diesel cut undergoes a hydrocracking or hydroisomerization step to obtain a hydrocracked effluent comprising a kerosene cut. Unlike the present invention, said process therefore comprises either a step of hydrocracking the diesel cut obtained, or a step of recycling said diesel cut in the hydroisomerization step.

[0021] In a third embodiment claimed in claim 19, said process comprises a step of separating said hydroisomerized effluent into a vapor effluent and a liquid effluent; a step of distilling the liquid effluent or the hydroisomerized effluent to obtain a kerosene cut and a diesel cut, and a step of hydrocracking said diesel cut to obtain a hydrocracked effluent comprising a kerosene cut. Unlike the present invention, said process therefore allows the co-production of diesel and kerosene cuts which are separated by a distillation step and the diesel cut obtained then undergoes a hydrocracking step.

[0022] US Patent 8,324,439 B2 teaches a process for treating renewable feedstocks of plant or animal origin. Said process comprises a step of hydrotreating the renewable feedstock, a step of separating the hydrotreated effluent to obtain hydrogen, other gases and at least one effluent containing hydrocarbons. Said process then comprises a step of hydroisomerizing at least a portion of said effluent containing hydrocarbons in the presence of a selective hydroisomerization catalyst, said catalyst comprising at least one one-dimensional zeolite with 10 MR and at least one metal from group VIII and / or group VIB. Finally, said process comprises a step of separating the hydroisomerized effluent to obtain hydrogen, other gases and at least one diesel-type cut.Said patent also teaches a process comprising a step of hydroisomerization of an effluent from the step of hydrotreatment of a renewable feedstock of animal or plant origin on a hydroisomerization catalyst. Said catalyst comprises at least one monodimensional zeolite with 10 MR of structural code TON or EUO, or a zeolite ZSM-48, ZBM-30, IZM-1, COK-7, EU-2 and EU-11, alone or in a mixture. The structural codes are defined in the classification of the International Zeolite Association (IZA: http: / / www.iza-structure.org / databases / ).

[0023] Patent application US 2014 / 0291200 A1 describes a process for producing diesel with a sulfur content of less than 10 ppm from renewable feedstocks, comprising the following steps: a) A step of hydrotreating renewable feedstocks b) A step of separating the hydrotreated effluent from a) into a hydrogen-rich gaseous fraction and a liquid hydrocarbon fraction c) A step of removing the H2S dissolved in the liquid hydrocarbon effluent from b) d) A step of hydroisomerizing at least a portion of the liquid effluent from step b) on a catalyst comprising at least one metal from group VIII or a combination of a metal from group VIII and a metal from group VI B on an amorphous support of alumina or silica-alumina type or crystalline, i.e. comprising a 10MR zeolite chosen from zeolites of structural type TON, EU-0 or chosen from zeolites ZSM-48, ZBM-30, IZM-1, COK-7, EU-2 and EU-11 alone or in mixture,e) A step of separating the resulting effluent into a hydrogen-rich gaseous fraction and at least one diesel fraction,

[0024] The recycling of gaseous fractions from separation steps b) and / or e) in steps a) and / or d) is mentioned with the aim of better controlling the exothermicity of the reactions in step a). No recycling of a part of the diesel fraction obtained or of a fraction heavier than diesel in steps a) and / or e) is mentioned.

[0025] Patent application US 2021 / 0395620 A1 describes a process for producing kerosene from renewable feedstocks. The process comprises a first step of hydrotreating the renewable feedstock in dilution with a hydrocarbon cut, followed by a hydrocracking step. The hydrocracked effluent is then separated into at least one gaseous fraction and at least one liquid fraction. More specifically, the separation step allows the production of a kerosene-type cut and a heavy diesel-type cut. At least a portion of the hydrocracked effluent, more specifically the kerosene-type cut, is then directed to the hydrodearomatization step in order to make it compatible with current product specifications, in particular the reduction of the aromatics content.

[0026] Patent applications WO 2010 / 000934 A1 and US 2010 / 000908 describe a process for producing kerosene from renewable feedstocks, comprising a first step of hydrodeoxygenation of renewable feedstocks allowing the production of n-paraffins. A second step according to the process allows the hydroisomerization of the n-paraffins from the first step and the production of an effluent rich in iso-paraffins. The latter is sent to a separation and fractionation step allowing the production of a cut boiling in the gasoline range, a cut boiling in the kerosene range and possibly a cut boiling in the diesel range and a cut having a boiling point greater than or equal to 200°C.The fraction having a boiling point greater than or equal to 200°C, optionally mixed with the fraction boiling in the diesel range, is then specifically recycled to a hydroisomerization step which may be identical to or distinct from the first hydroisomerization step in order to be reisomerized therein. The isomerization catalyst(s) comprise a group VIII metal and a support comprising a zeolite chosen from SAPO-11, SAPO-41, ZSM-22, ZSM-23 or ferrierite.

[0027] Patent application US 2014 / 0005450 A1 describes a method for manufacturing synthetic distillates, based on the hydrotreatment of a feedstock then its hydroisomerization / hydrocracking followed by a step of separating the paraffins (n-paraffins and i-paraffins) carried out on a molecular sieve. More specifically, the passage on a molecular sieve makes it possible to separate a cut rich in n-paraffins and a cut rich in i-paraffins. The cut rich in i-paraffins is then fractionated by distillation allowing the production of at least one fraction of heavy distillates and at least one fraction of light distillates or diesel. A recycle of the n-paraffins to the hydroisomerization / hydrocracking step is described. A recycle of a portion of the heavy distillates fraction is also described. This is a selective recycle of a stream according to its chemical family and not according to its distillation interval.The hydroisomerization catalyst used may comprise a group VIII metal selected from Pt and Pd, alone or in combination and a support which may be amorphous or crystalline selected from alumina, amorphous silica alumina, ferrierite, ALPO-31, SAPO-11, SAPO-31, SAPO, 37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-11, MeAPO-31, MeAPO-41, MgAPSO-11, MgAPSO-31, MgAPSO-41, MgAPSO-46, ELAPSO-11, ELAPSO-3, EMAPSO-41, ELAPSO-11, ELAPSO-31, ELAPSO-41, laumontite, cancrinite, offertite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.

[0028] Patent application US 2017 / 0022424 A1 describes a process for hydrotreating renewable feedstocks for producing an n-paraffinic hydrocarbon fraction. This patent claims a very specific process implementation of injecting renewable feedstock into different catalytic beds, mixed with a gas recycle stream (rich in hydrogen) and / or with a liquid recycle stream. The recycled streams (gas and liquid) are obtained from an effluent separation section positioned downstream of the hydrotreatment stage and are injected according to a ratio relative to the renewable feedstock advantageously selected to control the reaction exotherm and the temperature profile of the reactor.A portion of the liquid stream comprising the paraffins is sent to a hydroisomerization step in the presence of a hydroisomerization catalyst comprising at least one metal from group VIII chosen from platinum and palladium, nickel and cobalt and / or at least one metal from group VI B chosen from molybdenum and tungsten and at least one amorphous support of doped alumina, silica alumina type or a zeolitic support preferably comprising a zeolite of structural type TON (Nll-10), FER (ferrierite), EUO (EU-1 or ZSM-50) or the zeolites ZSM-48, ZBM-30, IZM-1, COK-7, EU-2 and Ell-11 alone or as a mixture. All or part of the hydroisomerized effluent is then sent to a fractionation stage to obtain a gaseous fraction, a naphtha fraction and a middle distillate fraction comprising diesel and kerosene. Part of the middle distillate fraction (150°C+) can be recycled to the hydrotreatment stage or to the hydroisomerization stage.In another variant, part of the 300°C+ fraction can be recycled in the hydroisomerization stage in order to upgrade this cut into lighter products and increase the cold properties.

[0029] Patent application US 2011 / 0105812 A1 describes a process for improving the cold properties of a diesel fraction produced from renewable feedstocks. The process comprises a step of hydrogenation / deoxygenation of the renewable feedstocks allowing the production of an n-paraffinic effluent followed by a step of hydroisomerization and selective hydrocracking of the n-paraffinic effluent. The effluent collected after hydroisomerization and selective hydrocracking is then selectively separated to produce at least one diesel stream comprising a portion of the paraffins having a boiling point in the diesel range, at least one recycled stream comprising another portion of the paraffins having a boiling point in the diesel range (preferably C16+), the composition of the diesel stream and the recycled stream being different, at least one light stream containing at least the LPGs and at least one naphtha-type stream.The recycled stream is advantageously recycled to the hydroisomerization and selective hydrocracking stage. The hydroisomerization catalyst used may comprise a group VIII metal selected from Pt and Pd, alone or in combination and a support which may be amorphous or crystalline selected from alumina, amorphous silica alumina, ferrierite, ALPO-31, SAPO-11, SAPO-31, SAPO, 37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-11, MeAPO-31, MeAPO-41, MgAPSO-11, MgAPSO-31, MgAPSO-41, MgAPSO-46, ELAPSO-11, ELAPSO-3, EMAPSO-41, ELAPSO-11, ELAPSO-31, ELAPSO-41, laumontite, cancrinite, offertite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.Patent application WO 2021 / 099343 A1 describes the composition of a hydrocarbon fraction rich in iso-paraffins and more specifically the composition of a hydrocarbon fraction having an excellent freezing point, compatible with an aviation fuel application. More particularly, the object of the invention is to provide an aviation fuel type hydrocarbon fraction with increased yield and density and produced from a renewable feedstock. Said composition is obtained by a process comprising the hydrodeoxygenation of a renewable feedstock of vegetable oil type followed by a step of isomerization of the n-paraffins formed to produce iso-paraffins, the isomerization catalyst comprising a group VIII metal chosen from platinum, palladium, and nickel and a support comprising a zeolite chosen from SAPO-11, SAPO-41, ZSM-22, ZSM-23 or ferrierite.The iso-paraffins formed can then be fractionated to obtain said composition.

[0030] For the purposes of the present invention, the various embodiments presented can be used alone or in combination with each other, without limitation of combination.

[0031] For the purposes of the present invention, the different parameter ranges for a given step such as pressure ranges and temperature ranges may be used alone or in combination. For example, for the purposes of the present invention, a preferred range of pressure values ​​may be combined with a more preferred range of temperature values.

[0032] In the following text, the term naphtha cut or naphtha is understood to mean the hydrocarbon fraction having a boiling point lower than the middle distillate cut. The middle distillate cut generally has an initial cutting point between 120°C and 160°C, preferably 120°C. The naphtha cut can have boiling points ranging from that of hydrocarbon compounds having 5 carbon atoms per molecule (or 36°C boiling point) up to 216°C and includes the gasoline cut.

[0033] Throughout the remainder of the text, the term kerosene or kerosene cut means the cut having initial and final boiling points between 120°C and 300°C and the term diesel or gas oil cut means a cut having initial and final boiling points in a range from 120 to 400°C and preferably between 120 and 380°C.

[0034] In the following text, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals in columns 8, 9 and 10 according to the new IIIPAC classification, and group VI B to the metals in column 6.

[0035] In the remainder of the text, the expressions "between ... and ..." and "between .... and ..." are equivalent and mean that the limit values ​​of the interval are included in the range of values ​​described. If this were not the case and the limit values ​​were not included in the range described, such precision will be provided by the present invention.

[0036] In this description, the expression "greater than..." is understood as strictly greater than, and symbolized by the sign ">", and the expression "less than" as strictly less than, and symbolized by the sign "<".

[0037] ADVANTAGES OF THE METHOD ACCORDING TO THE INVENTION

[0038] In attempting to develop a process for treating a feedstock from a renewable source for the selective production of a kerosene cut, the applicant discovered that the use of a specific bi-functional catalyst comprising a hydrogenating / dehydrogenating phase based on a noble metal (Pt) and a support comprising a specific zeolite chosen from zeolites with structural code MTW, and zeolite IZM-2 alone or as a mixture for the hydroconversion step (hydroisomerization and / or hydrocracking reactions) of linear paraffins from the hydrodeoxygenation step of renewable feedstocks, combined with the implementation of a recycling step in the hydroconversion step, of all or part of a heavy cut having an initial boiling point of between 250 and 300°C separated at the end of the hydroconversion step, was of great interest.

[0039] Subject of the invention

[0040] More specifically, the present invention relates to a process for treating a feedstock from a renewable source to produce a kerosene cut, said process comprising at least the following steps, and preferably consisting of the following steps: a) a step of hydrotreating said feedstock in the presence of a fixed-bed catalyst, said catalyst comprising a hydrogenating function and an oxide support, at a temperature of between 200 and 450°C, at a pressure of between 1 MPa and 10 MPa, at an hourly space velocity of between 0.1 h -1 and 10 a.m. -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 2000 Nm 3 hydrogen / m 3of charge, b) a step of separating at least a portion of the effluent from step a) into at least one light gaseous fraction, at least one hydrocarbon liquid effluent, and at least one aqueous liquid effluent, c) a step of hydroconversion of at least a portion of the hydrocarbon liquid effluent from step b) in the presence of at least one fixed-bed bifunctional hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one noble metal from group VIII of the periodic table chosen from platinum and palladium and a support comprising at least one zeolite chosen from zeolites with structural code MTW, and zeolite IZM-2 alone or as a mixture, and at least one binder, said hydroconversion step being carried out at a temperature of between 250°C and 500°C, at a pressure of between 1 MPa and 10 MPa, at an hourly space velocity of between 0.1 and 10 a.m. -1and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3 of charge, d) a step of separating at least a portion of the effluent from step c) which makes it possible to separate at least one gaseous fraction, and at least one liquid hydrocarbon effluent, e) a step of fractionating the hydrocarbon effluent from step d) into at least one kerosene fraction, at least one heavy fraction having an initial boiling point of between 250 and 300°C,

[0041] Said process comprising a step of recycling all or part of said heavy cut having an initial boiling point of between 250 and 300°C from said fractionation step e) in said hydroconversion step c).

[0042] An advantage of the present invention is to provide a method for treating a feedstock from a renewable source in two stages to produce a kerosene cut, using a specific catalyst in the hydroconversion stage making it possible to obtain a gain in activity and selectivity. The use of the specific catalyst makes it possible, all other things being equal, to reduce the temperature necessary to obtain the cold property targets on the kerosene cut produced. The implementation of the hydroconversion stage with the specific catalyst according to the invention also makes it possible, for the cold property targets aimed at, to maximize the yield of said kerosene cut of interest produced in said process.Another advantage of the present invention is to provide a process for treating a feedstock from a renewable source to produce at least one kerosene cut, in 2 stages and comprising a hydrotreatment stage and a hydroconversion stage making it possible to obtain a gain in kerosene cut yield by implementing a specific catalyst in the hydroconversion stage and recycling in said hydroconversion stage a heavy cut having an initial boiling point of between 250 and 300°C.

[0043] The implementation of the process and in particular of such recycling makes it possible to maximize the selectivity of the process towards said kerosene cut of interest while maximizing the conversion of the renewable feedstock. More particularly, the operating conditions of the hydroconversion step as well as the cutting point of said kerosene cut can be jointly adapted to maximize the yields in the kerosene cut of interest.

[0044] Another advantage of the present invention is to provide a method of fractionating the hydroconversion effluent which offers flexibility in adjusting the distillation intervals of the kerosene cut (cut point advantageously adjustable in the range 120 to 300°C for maximum kerosene production).

[0045] An advantage of the present invention is to allow the treatment of heavier renewable source feedstocks within said process thanks to the implementation of step e) of fractionation and recycling of said heavy cut in step c) of hydroconversion compared to processes of the prior art not implementing such a fractionation step.

[0046] Detailed description of the invention

[0047] Charges

[0048] The present invention is particularly dedicated to the preparation of diesel and / or kerosene fuel bases corresponding to new environmental standards, from charges from renewable sources.

[0049] The feedstocks from renewable sources used in the process according to the present invention are advantageously chosen from vegetable oils (for example palm, rapeseed, soybean), animal fats, used cooking oils, oils of microbial origin (for example from algae), fish oils, long paraffins (waxes) from the Fischer-Tropsch process, crude or having undergone a pretreatment, or mixtures of such feedstocks, containing triglycerides and / or free fatty acids and / or esters. The vegetable oils may advantageously be crude or refined, totally or in part, and derived from the following plants: rapeseed, sunflower, soybean, palm, palm kernel, olive, coconut, jatropha, this list not being exhaustive. Algal or fish oils are also relevant.Animal fats are advantageously chosen from lard or fats composed of residues from the food industry or from the catering industry.

[0050] These fillers essentially contain triglyceride-type chemical structures that the skilled person also knows as fatty acid triester as well as free fatty acids. A fatty acid triester is thus composed of three fatty acid chains. These fatty acid chains in the form of triester or in the form of free fatty acids, have a number of unsaturations per chain, also called the number of carbon-carbon double bonds per chain, generally between 0 and 3 but which can be higher, particularly for oils derived from algae which generally have a number of unsaturations per chain of 5 to 6.

[0051] The molecules present in the fillers from renewable sources used in the present invention therefore have a number of unsaturations, expressed per triglyceride molecule, advantageously between 0 and 18. In these fillers, the level of unsaturation, expressed as the number of unsaturations per hydrocarbon fatty chain, is advantageously between 0 and 6.

[0052] Feedstocks from renewable sources generally also contain various impurities, including heteroatoms such as nitrogen. Nitrogen levels in vegetable oils are generally between approximately 1 ppm and 1000 ppm by weight, depending on their nature.

[0053] Process and catalysts

[0054] Advantageously, the feedstock may undergo, prior to step a) of the process according to the invention, a pre-treatment or pre-refining step so as to eliminate, by appropriate treatment, contaminants such as metals, such as alkaline compounds, for example on ion exchange resins, alkaline earths and phosphorus. Suitable treatments may, for example, be thermal and / or chemical treatments well known to those skilled in the art.

[0055] In accordance with step a) of the process according to the invention, the feedstock, optionally pretreated, is brought into contact with a fixed-bed catalyst at a temperature of between 200 and 450°C, preferably between 220 and 350°C, preferably between 220 and 320°C, and even more preferably between 220 and 310°C. The pressure is between 1 MPa and 10 MPa, preferably between 1 MPa and 6 MPa and even more preferably between 1 MPa and 4 MPa. The hourly space velocity, i.e. the volume of feedstock per volume of catalyst per hour, is between 0.1 h -1 and 10 a.m. 1 The charge is brought into contact with the catalyst in the presence of hydrogen. The total quantity of hydrogen mixed with the charge is such that the hydrogen / charge ratio is between 70 and 2000 Nm 3 hydrogen / m 3 load and preferably between 150 and 1000 Nm 3 hydrogen / m 3 dump.

[0056] In step a) of the process according to the invention, the fixed-bed catalyst is advantageously a hydrotreatment catalyst comprising a hydro-dehydrogenating function comprising at least one metal from group VIII and / or group VI B, taken alone or as a mixture and a support chosen from the group formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also advantageously contain other compounds and for example oxides chosen from the group formed by boron oxide, zirconia, titanium oxide, phosphoric anhydride. The preferred support is an alumina support and very preferably r], 5 or y alumina.

[0057] Said catalyst is advantageously a catalyst comprising metals from group VIII preferably chosen from nickel and cobalt, taken alone or in a mixture, preferably in association with at least one metal from group VI B preferably chosen from molybdenum and tungsten, taken alone or in a mixture.

[0058] The content of metal oxides of groups VIII and preferably nickel oxide is advantageously between 0.5 and 10% by weight of nickel oxide (NiO) and preferably between 1 and 5% by weight of nickel oxide and the content of metal oxides of groups VI B and preferably molybdenum trioxide is advantageously between 1 and 30% by weight of molybdenum trioxide (MoOa), preferably 5 to 25% by weight, the percentages being expressed in % by weight relative to the total mass of the catalyst. The total content of metal oxides of groups VI B and VIII in the catalyst used in step a) is advantageously between 5 and 40% by weight and preferably between 6 and 30% by weight relative to the total mass of the catalyst.

[0059] Said catalyst used in step a) of the process according to the invention must advantageously be characterized by a high hydrogenating power so as to orient the selectivity of the reaction as much as possible towards a hydrogenation conserving the number of carbon atoms of the fatty chains, i.e. the hydrodeoxygenation route, in order to maximize the yield of hydrocarbons entering the distillation field of kerosenes and / or diesels. This is why it is preferred to operate at a relatively low temperature. Maximizing the hydrogenating function also makes it possible to limit the polymerization and / or condensation reactions leading to the formation of coke which would degrade the stability of the catalytic performances. Preferably, a Ni or NiMo type catalyst is used.

[0060] Said catalyst used in hydrotreatment step a) of the process according to the invention may also advantageously contain a doping element chosen from phosphorus and boron, taken alone or in a mixture. Said doping element may be introduced into the matrix or preferably be deposited on the support. Silicon may also be deposited on the support, alone or with phosphorus and / or boron and / or fluorine.

[0061] The weight content of oxide of said doping element is advantageously less than 20% by weight and preferably less than 10% by weight and it is advantageously at least 0.001% by weight.

[0062] Preferred catalysts are the catalysts described in patent application FR 2 943 071 describing catalysts having high selectivity for hydrodeoxygenation reactions.

[0063] Other preferred catalysts are the catalysts described in patent application EP 2 210 663 describing supported or bulk catalysts comprising an active phase consisting of a sulfur-containing element from group VIB, the element from group VIB being molybdenum.

[0064] The metals of the catalysts used in hydrotreatment step a) of the process according to the invention are sulphide metals or metallic phases and preferably sulphide metals. It would not be outside the scope of the present invention to use in step a) of the process according to the invention, simultaneously or successively, a single catalyst or several different catalysts. This step can be carried out industrially in one or more reactors with one or more catalytic beds and preferably with a descending liquid flow.

[0065] Said hydrotreatment step a) allows the hydrogenation, hydrodeoxygenation, hydrodenitrogenation and hydrodesulfurization of said feed.

[0066] In accordance with step b) of the process according to the invention, a step of separating at least a portion and preferably all of the effluent from step a) is carried out. Said step b) makes it possible to separate at least one so-called light gaseous fraction rich in hydrogen, at least one liquid hydrocarbon effluent consisting of n-paraffins, and at least one aqueous liquid effluent.

[0067] Said light gaseous fraction comprises at least the hydrogen not converted by the reactions carried out in step a), at least the gases with one or more oxygen atoms resulting from the decomposition of the oxygenated compounds in step a) and at least the C4 compounds; that is to say the compounds C1 to C4 preferably having a final boiling point of less than 20°C. The aim of this step is to separate the gases from the liquids. More particularly, the aim is to recover at least the hydrogen-rich gases which may also contain compounds such as CO and CO2, at least one liquid hydrocarbon effluent consisting of n-paraffins and at least one aqueous liquid effluent containing the water produced by the reactions carried out in step a). Said liquid hydrocarbon effluent preferably has a sulfur content of less than 10 ppm by weight, a nitrogen content of less than 2 ppm by weight.

[0068] Separation step b) can advantageously be implemented by any method known to those skilled in the art, such as, for example, the combination of one or more high and / or low pressure separators operated hot or cold, and / or distillation steps, and / or high pressure and / or low pressure stripping.

[0069] Step b) also allows the separation of at least one aqueous liquid effluent, preferably water. The removal of at least part of the water and preferably all of the water can be carried out by any methods and techniques known to those skilled in the art. Preferably, the removal of the water is carried out by decantation in a separator drum or by drying or by passing through a desiccant or by flash or by a combination of at least two of these techniques. The atomic oxygen content of the liquid hydrocarbon effluent containing the paraffinic hydrocarbons resulting from step b) of the process according to the invention, expressed in parts per million by weight (ppm), is preferably less than 10,000 ppm, preferably less than 6,000 ppm, very preferably less than 1,000 ppm by weight, even more preferably less than 500 ppm by weight.The content in ppm by weight of atomic oxygen in said hydrocarbon liquid effluent is measured by the infrared absorption technique such as for example the technique described in patent application US2009 / 0018374A1.

[0070] According to the invention, the process comprises a step c) of hydroconversion of at least a portion and preferably all of the hydrocarbon liquid effluent from step b) of the process in the presence of at least one fixed-bed bifunctional hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one noble metal from group VIII of the periodic table chosen from platinum and palladium alone or in a mixture and a support comprising at least one zeolite chosen from zeolites with structural code MTW and IZM-2 alone or in a mixture and at least one binder, said hydroconversion step being carried out at a temperature of between 250 and 500°C, at a pressure of between 1 MPa and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3load, and preferably between 150 and 750 Nm 3 / m 3 dump.

[0071] The operating conditions of hydroconversion step c) are adjusted to promote the hydroisomerization or hydrocracking reactions as required. Preferably, hydroconversion step c) of the process according to the invention is a hydroisomerization step and advantageously operates at a temperature of between 250°C and 450°C, and very preferably between 250 and 400°C, at a pressure of between 2 MPa and 10 MPa and very preferably between 3 MPa and 9 MPa, at an hourly volumetric flow rate advantageously of between 0.2 and 7 h -1 and very preferably, between 0.5 and 5 h -1 (defined in relation to the charge flow rate of step c), excluding recycle) at a hydrogen flow rate such that the hydrogen / charge volume ratio is advantageously between 100 and 1000 normal m 3 of hydrogen per m 3load and preferably between 150 and 1000 normal m 3 of hydrogen per m 3 of charge (ratio defined in relation to the charge flow rate of step c), excluding recycle). The process according to the invention is a flexible process which allows, by the implementation of adjusted operating conditions in the hydroconversion step c), combined with the implementation of a specific catalyst in said step c) and a final fractionation step e) allowing the separation of at least one heavy cut having an initial boiling point of between 250 and 300°C which is recycled in said step c), to maximize the yield of kerosene cut obtained.

[0072] According to the invention, part and preferably all of the liquid hydrocarbon effluent from step b) is sent to said hydroconversion step c) in a mixture with all or part of the heavy cut having an initial boiling point of between 250 and 300°C which is recycled in said step c) and resulting from the final fractionation step e).

[0073] Said hydroconversion step may advantageously comprise one or more catalytic beds which may comprise one or more different catalysts in each of the beds.

[0074] According to a preferred embodiment, the heavy fraction having an initial boiling point of between 250 and 300°C from fractionation step e) is advantageously mixed with all or part of the liquid hydrocarbon effluent from step b) upstream of the hydroconversion reactor and then this mixture is injected into the first catalytic bed of step c) in the case where said hydroconversion step comprises several catalytic beds.

[0075] In another embodiment, the heavy fraction having an initial boiling point of between 250 and 300°C from fractionation step e) is advantageously injected into an intermediate bed of the hydroconversion reactor in the case where said hydroconversion step comprises several catalytic beds. In this optional embodiment, the beds preceding the recycle injection then only treat all or part of the liquid hydrocarbon effluent from step b) of the process according to the invention.

[0076] In all cases, the operating conditions of hydroconversion step c) are adjusted to promote the hydroisomerization or hydrocracking reactions as required. More particularly, the operating conditions of hydroconversion step c) and the recycle rate of the heavy cut having an initial boiling point of between 250 and 300°C from fractionation step e) may be advantageously adjusted in order to maximize the yields of the desired kerosene cut at the outlet of said process. In accordance with the invention, the catalyst used in step c) is a bifunctional catalyst comprising a hydrogenating phase comprising at least one noble metal from group VIII of the periodic table chosen from platinum and palladium alone or as a mixture and a support comprising at least one zeolite chosen from zeolites of MTW structural type, and IZM-2 zeolite alone or as a mixture and at least one binder.

[0077] The hydro / dehydrogenating function

[0078] Preferably, the group VIII metal of the catalyst used in step c) is platinum. Advantageously, the hydro / dehydrogenating (metallic) element and preferably platinum can be introduced onto the catalyst support by any method known to those skilled in the art, such as, for example, co-mixing, dry impregnation, or exchange impregnation. According to one or more embodiments, the group VIII metal content, and preferably the platinum content, in the catalyst used in step c) is between 0.01% by weight and 4% by weight, preferably between 0.05% by weight and 2% by weight, relative to the total weight of said catalyst.

[0079] The catalyst used in step c) may also advantageously further comprise at least one additional metal chosen from the group formed by the metals of groups II IA, IVA and VI IB of the periodic table of elements and preferably chosen from gallium, indium, tin and rhenium. Said additional metal is preferably chosen from indium, tin and rhenium.

[0080] Preferably, the content of at least one additional metal in the catalyst used in step c) is between 0.01% by weight and 2% by weight, preferably between 0.05% by weight and 1% by weight, relative to the total weight of said catalyst.

[0081] According to one embodiment, the sulfur content in the hydroconversion catalyst is such that the ratio of the number of moles of sulfur to the number of moles of the at least one metal of group VI 11 B is between 0.3 and 3. According to one or more embodiments, the presence of sulfur in the catalyst comes from an optional sulfurization step of the hydroconversion catalyst. According to one or more embodiments, the presence of sulfur in the catalyst comes from potentially present impurities, such as for example in the alumina binder. According to another embodiment, the catalyst does not contain sulfur.

[0082] The acid function. According to the invention, the catalyst used in step c) comprises a support comprising at least one zeolite chosen from zeolites of MTW structural type, and IZM-2 zeolite alone or as a mixture and at least one binder.

[0083] The acid function of the bifunctional catalyst used in step c) is provided by the zeolite chosen from MTW structural type zeolites, and IZM-2 zeolite, alone or in a mixture.

[0084] Preferably, the MTW structural type zeolites are chosen from the ZSM-12, TPZ-12, Theta-3, Nll-13, CZH-5 zeolites, alone or as a mixture and preferably, the MTW structural type zeolite is ZSM-12.

[0085] According to a preferred embodiment, the catalyst used in step c) comprises a support comprising an IZM-2 zeolite alone or a support comprising a ZSM-12 zeolite alone. The IZM-2 zeolite is a crystallized microporous solid whose crystal structure and preparation process are described in patent application FR2918050A1. The structural code of the IZM-2 zeolite is not known to date. The ZSM-12 zeolite is a crystallized microporous solid whose crystal structure is described on the website of the International Zeolyst Association (http: / / www.iza-structure.org / ). It is a one-dimensional zeolite with 12 MR, its structural code is MTW. A process for preparing this zeolite is for example described in the article Synthesis of zeolite ZSM-12 in the system (MTEA)2O-Na2O-SiO2-Al2O3-H2O by S. Ersnt et al. (Zeolites, 7, 5, 458-462, DOI10.1016 / 0144-2449(87)90015-7).

[0086] The zeolites are preferably essentially in acid form, that is to say that the atomic ratio between the monovalent compensation cation (for example sodium) and the aluminum inserted in the crystal lattice of the solid is advantageously less than 0.1, preferably less than 0.05 and very preferably less than 0.01. According to one or more embodiments, the zeolites entering into the composition of said hydroisomerization catalyst are advantageously calcined. According to one or more embodiments, said zeolites are exchanged by at least one treatment with a solution of at least one ammonium salt so as to obtain the ammonium form of the zeolites which, once calcined, leads to the acid form of said zeolites.

[0087] In a preferred embodiment, the catalyst used in step c) comprises a hydrogenating phase comprising platinum and a support comprising an IZM-2 zeolite alone and an alumina binder. In another preferred embodiment, the catalyst used in step c) comprises a hydrogenating phase comprising platinum and a support comprising an ZSM-12 zeolite alone and an alumina binder.

[0088] Preferably, the catalyst used in step c) comprises a zeolite content of between 1% by weight and 90% by weight, preferably between 3% by weight and 80% by weight, and more preferably between 4% by weight and 60% by weight, preferably between 4 and 30% by weight and more preferably between 4 and 20% by weight relative to the total weight of said catalyst.

[0089] The binder

[0090] Preferably, the binder of the catalyst support of step c) is amorphous or crystallized. Preferably, the binder used in the catalyst support of step c) is advantageously chosen from the group formed by alumina, silica, silica-alumina, clays, titanium oxide, boron oxide, zirconia and aluminates, taken alone or as a mixture. Preferably, the binder is alumina. Preferably, said binder may contain alumina in all its forms known to those skilled in the art, such as for example alpha, gamma, eta, delta type aluminas.

[0091] Preferably, the catalyst used in step c) comprises a binder content of between 10% by weight and 99% by weight, relative to the total weight of said catalyst, i.e., so as to ensure the addition to 100% by weight of the elements constituting the catalyst used in step c).

[0092] According to the invention, the catalyst support comprises the zeolite mixed with a binder. The shaping of the support in the form of a mixture is preferably carried out by co-mixing, extrusion and then heat treatment of the zeolite with the binder or a precursor of the binder, such as for example boehmite, which by heat treatment is transformed into alumina.

[0093] A preferred catalyst of step c) comprises and preferably consists of platinum, and a support comprising and preferably consists of a ZSM-12 zeolite and an alumina binder. Another preferred catalyst of step c) comprises and preferably consists of platinum, and a support comprising and preferably consists of an IZM-2 zeolite and an alumina binder.

[0094] According to a preferred embodiment, the catalyst used in step c) more particularly comprises, and preferably consists of:

[0095] - from 1% to 90% by weight, preferably from 3% to 80% by weight and even more preferably from 4% to 60% by weight of zeolite;

[0096] - from 0.01% to 4% by weight, preferably from 0.05% to 2% by weight of at least one metal from group VI II B, preferably platinum;

[0097] - optionally from 0.01% to 2% by weight, preferably from 0.05% to 1% by weight of at least one additional metal chosen from the group formed by the metals of groups 11 IA, IA and VII B;

[0098] - optionally a sulfur content, preferably such that the ratio of the number of moles of sulfur to the number of moles of metal(s) of group VI 11 B is between 0.3 and 3; and

[0099] - optionally at least one binder, preferably alumina, ensuring the addition to 100% by weight in the catalyst, relative to the total weight of the catalyst from step c).

[0100] Preferably, the catalyst used in step c) is shaped in the form of cylindrical or polylobed extrudates such as bilobed, trilobed, polylobed of straight or twisted shape. According to one or more embodiments, the catalyst used in step c) is shaped in the form of crushed powders, tablets, rings, balls, wheels. Techniques other than extrusion, such as pelletizing or coating, can advantageously be used.

[0101] Preferably, the noble metal contained in said catalyst used in step c) can advantageously be reduced. One of the preferred methods for carrying out the reduction of the metal is treatment at a temperature between 150°C and 650°C and a total pressure between 0.1 and 25 MPa. For example, a reduction may comprise a two-hour stage at 150°C followed by a temperature increase to 450°C at a rate of 1 oC / min then a two-hour stage at 450°C; during the reduction stage, the hydrogen flow rate can be 1000 normal m 3 hydrogen / m 3 catalyst and the total pressure can be kept constant at 0.1 MPa. Any reduction method can advantageously be considered, either in situ (the reduction of the catalyst is carried out in the same unit where the catalytic reaction is carried out), or ex situ (the reduction is carried out outside the unit where the catalytic reaction is carried out, before loading the catalyst into the unit).

[0102] According to the invention, the process comprises a step d) of separating at least part and preferably all of the effluent from step c).

[0103] Said step d) makes it possible to separate at least one so-called light gaseous fraction and at least one hydrocarbon liquid effluent. Optionally, step d) also makes it possible to separate at least part of the residual water and preferably all of the residual water.

[0104] Said so-called light gaseous fraction comprises at least the hydrogen not converted by the reactions described in step c) and at least a portion of the C4 cracking products; i.e. compounds C1 to C4 preferably having a final boiling point below 20°C. The purpose of this step is to separate the gases from the liquids. More particularly, the purpose is to recover at least the hydrogen-rich gases, at least one liquid hydrocarbon effluent rich in branched paraffins and optionally an aqueous liquid effluent containing a small amount of residual water (residual water generated either by the hydrodeoxygenation of the residual oxygenated compounds present in the feedstock of step c) or by a slight entrainment of water from step b) to step c)).

[0105] The separation step d) described can advantageously be implemented by any method known to those skilled in the art such as, for example, the combination of one or more high and / or low pressure separator flasks operated hot or cold, and / or distillation steps, and / or high pressure and / or low pressure stripping.

[0106] In accordance with the process according to the invention, the liquid hydrocarbon effluent from step d) undergoes a step e) of fractionation into at least one kerosene fraction, at least one heavy fraction having an initial boiling point of between 250 and 300°C and preferably between 270 and 300°C.

[0107] Advantageously, said fractionation step e) also allows the separation of at least one gaseous fraction consisting mainly of light C1-C4 hydrocarbons, and at least one lighter hydrocarbon fraction called naphtha.

[0108] Fractionation step e) may be carried out by any method known to those skilled in the art and preferably may be carried out in a distillation column or in a steam stripping step followed by a distillation column. Said heavy cut having an initial boiling point of between 250 and 300°C separated in said step e) is advantageously mainly composed of unconverted long n-paraffins, i.e. non-isomerized and non-cracked in step c) and is advantageously obtained at the bottom of the distillation column.

[0109] More particularly, the cut point between the kerosene and naphtha cut is adjustable in the range between 100 and 150°C, and preferably around 120°C. More particularly, the cut point between the kerosene cut and the heavy cut is adjustable in the range between 250 and 300°C if the aim is to target the production of kerosene, and preferably around 300°C.

[0110] The naphtha cut thus contains compounds which have a boiling point below 120°C, the kerosene cut contains compounds which have a boiling point between 120°C and 300°C.

[0111] More particularly, the fractionation according to step e) of the process advantageously makes it possible to treat atypical feedstocks which generate, for example, long n-paraffins (number of carbon atoms greater than C20) within the hydroconversion step c). This is the case, for example, for n-paraffins derived from waxes produced by the Fischer-Tropsch process: said feedstocks can advantageously be treated in a mixture with the vegetable oil type feedstocks in the process according to the invention.

[0112] In this case, the specific hydroconversion catalyst used in step c) combined with the fractionation step e) of the process according to the invention advantageously makes it possible to maximize, on the one hand, the conversion of the long paraffins contained in the Fischer-Tropsch waxes and, on the other hand, to maximize the selectivity of the process towards the product of interest, which constitutes the kerosene cut.

[0113] The recovery of the bio-naphtha cut is not the subject of the present invention, but this cut can advantageously be sent to a steam cracking or reforming unit or even recovered as a gasoline base in a mixture with other gasoline bases.

[0114] According to the invention, said process comprises a step of recycling all or part of said unconverted heavy fraction from said fractionation step e) in said hydroisomerization step c). The process according to the invention advantageously does not comprise a hydrocracking step downstream of step c) of hydroconversion of part or all of the effluent from step c).

[0115] Description of figures

[0116] [Fig 1] Figure 1 represents the different stages of the process for producing kerosene or renewable diesel in 2 stages with recycle, using a specific catalyst for stage c) of hydroconversion.

[0117] The feedstock from renewable sources is sent via line 1 mixed with make-up and / or recycled hydrogen (line 2) to a hydrotreatment step a) (enabling hydrogenation, hydrodeoxygenation, hydrodenitrogenation and hydrodesulfurization of the feedstock). The hydrotreated effluent from hydrotreatment step a) is withdrawn via line 3 and is sent to three-phase separation step b) which separates at least one hydrogen-rich gaseous effluent (line 4) and at least one hydrocarbon liquid effluent (line 5). Step b) also removes at least part of the water produced by the hydrodeoxygenation reactions and preferably all of this water (line 6).

[0118] The liquid hydrocarbon effluent (5) is sent to a hydroconversion step c) (hydroisomerization and / or hydrocracking), in the presence of a make-up and / or recycled hydrogen stream (line 7) and a heavy cut having an initial boiling point of between 250 and 300°C, which is recycled (line 15) from the fractionation step e) to produce a second effluent (line 8) which is sent to a three-phase separation step d) making it possible to separate a hydrogen-rich gaseous effluent (9) which may also contain light products such as the C1 - C4 cut, and at least one liquid hydrocarbon effluent (10). Said step also makes it possible to eliminate at least part of the water and preferably all of the residual water (11).

[0119] The liquid hydrocarbon effluent from step d) (10) is sent to a step e) of fractionation of the hydrocarbon effluent from step d) by distillation allowing the separation of a light gaseous fraction (12), a naphtha hydrocarbon cut (13), and a kerosene cut (14) and a heavy cut having an initial boiling point of between 250 and 300°C (15) which is recycled in the hydroconversion step c). Optionally, the fractionation allows the co-production of a kerosene cut (14) with a diesel cut (15).

[0120] The examples below illustrate the invention without limiting its scope.

[0121] EXAMPLES

[0122] Example 1: preparation of a hydrotreatment catalyst (C1).

[0123] The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with molybdenum oxide MoOa contents of 22% by weight, nickel oxide NiO of 4% by weight and phosphorus oxide P2O5 of 5% by weight relative to the total weight of the finished catalyst.

[0124] Example 2: Preparation of a hydroconversion catalyst in accordance with the invention (C2).

[0125] Synthesis of IZM-2 zeolite.

[0126] The IZM-2 zeolite was synthesized in accordance with the teaching of patent FR 2 918 050 B. A colloidal suspension of silica known under the commercial term Ludox HS-40 marketed by Aldrich, is incorporated into a solution composed of sodium hydroxide (Prolabo), structuring agent dibromide of 1,6bis(methylpiperidinium)hexane, aluminum hydroxide (Aldrich) and deionized water. The molar composition of the mixture is as follows: 1 SiCh; 0.0060 AI2O3; 0.1666 Na2O; 0.1666 1,6bis(methylpiperidinium)hexane; 33.3333 H2O. The mixture is stirred vigorously for half an hour. The mixture is then transferred, after homogenization, into a PARR type autoclave. The autoclave is heated for 5 days at 170°C with stirring on a rotating spit (30 rpm). The product obtained is filtered, washed with deionized water to reach a neutral pH and then dried overnight at 100°C in an oven. The solid is then introduced into a muffle furnace to be calcined to remove the structuring agent.The calcination cycle includes a temperature rise to 200°C, a two-hour hold at this temperature, a temperature rise to 550°C followed by an eight-hour hold at this temperature and finally a return to room temperature. The temperature rises are carried out with a ramp of 2°C / min. The solid thus obtained is then refluxed for 2 hours in an aqueous solution of ammonium nitrate (10 ml of solution per gram of solid, ammonium nitrate concentration of 3 M) in order to exchange the sodium alkali cations with ammonium ions. This refluxing step is carried out four times with a fresh solution of ammonium nitrate, then the solid is filtered, washed with deionized water and dried in an oven overnight at 100°C. Finally, to obtain the zeolite in its acid form (protonated H. +) a calcination step is carried out at 550°C for ten hours (temperature rise ramp of 2°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The solid thus obtained was analyzed by X-ray diffraction and identified as consisting of IZM-2 zeolite. Characterizations using NMR methods of 27 AI, X-ray fluorescence and ICP provide access to the following results for IZM-2:

[0127] - weight percentage of hexacoordinated aluminum atoms AI VI : 5%,

[0128] - ratio of the number of moles of silicon divided by the number of moles of aluminum in the network, in mole / mole, Si / Al: 72,

[0129] - ratio of the number of moles of sodium divided by the number of moles of network aluminum, in mole / mole, Na / AI: 0.03.

[0130] Preparation of IZM-2 / alumina support.

[0131] The IZM-2 / alumina support is obtained by mixing and extruding the IZM-2 zeolite with a Pural SB3 alumina gel. The mixed paste is extruded through a 1.8 mm diameter four-lobe die. After drying in an oven overnight at 110°C, the extrudates are calcined at 500°C for two hours (temperature rise ramp of 5°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The weight content of the IZM-2 zeolite in the support after calcination is 13% by weight.

[0132] Impregnation of platinum on IZM-2 / alumina support.

[0133] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NC>3)2. 50 grams of support are used and dry impregnated in a drageoir. After impregnation, the solid is left to mature for at least five hours in laboratory air and then left to dry overnight in an oven at 110°C. Finally, a calcination step is carried out under dry air flow (2 normal liters per hour and per gram of solid) in a tubular furnace under the following conditions:

[0134] - increase in temperature from ambient to 450°C at 5°C / min;

[0135] - two-hour stage at 450°C; - lowering to ambient temperature.

[0136] The Pt content measured by FX on the calcined C2 catalyst is 0.29% by weight, its dispersion measured by H2 / O2 titration is 52%, its distribution coefficient measured by Castaing microprobe is 0.91.

[0137] Example 3: Preparation of a hydroconversion catalyst in accordance with the invention (C3).

[0138] Zeolite ZSM-12.

[0139] The ZSM-12 zeolite is supplied by Zeolyst. The solid was analyzed by X-ray diffraction and identified as consisting of ZSM-12 zeolite. Characterizations using NMR methods of 27 AI, X-ray fluorescence and ICP provide access to the following results for ZSM-12:

[0140] - weight percentage of hexacoordinated aluminum atoms AI VI : 0% ;

[0141] - ratio of the number of moles of silicon divided by the number of moles of lattice aluminum, in mole / mole, Si / Al: 43;

[0142] - ratio of the number of moles of sodium divided by the number of moles of lattice aluminum, in mole / mole, Na / AI: 0.009.

[0143] Preparation of ZSM-12 / alumina support.

[0144] The ZSM-12 / alumina support is obtained by mixing and extruding the ZSM-12 zeolite with a Pural SB3 alumina gel. The mixed paste is extruded through a 1.8 mm diameter four-lobe die. After drying in an oven overnight at 110°C, the extrudates are calcined at 500°C for two hours (temperature rise ramp of 5°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The weight content of the ZSM-12 zeolite in the support after calcination is 6% by weight.

[0145] Impregnation of platinum on ZSM-12 / alumina support.

[0146] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NC>3)2. 50 grams of support are used and dry impregnated in a drageoir. After impregnation, the solid is left to mature for at least five hours in laboratory air and then left to dry overnight in an oven at 110°C. Finally, a calcination step is carried out under dry air flow (2 normal liters per hour and per gram of solid) in a tubular furnace under the following conditions:

[0147] - increase in temperature from ambient to 450°C at 5°C / min;

[0148] - two-hour stage at 450°C; - lowering to ambient temperature.

[0149] The Pt content measured by FX on the calcined C3 catalyst is 0.31% by weight, its dispersion measured by H2 / O2 titration is 38%, its distribution coefficient measured by Castaing microprobe is 0.87.

[0150] Example 4: hydrotreatment of a feedstock from a renewable source using a process in accordance with the invention

[0151] In a reactor regulated in temperature so as to ensure isothermal operation and with a fixed bed loaded with 190 ml of hydrotreatment catalyst C1, the catalyst being previously sulfurized, the hydrotreatment of pre-refined rapeseed oil with a density of 920 kg / m is carried out. 3 having an oxygen content of 11% by weight. The cetane index is 35 and the fatty acid distribution of rapeseed oil is detailed in Table 1. Prior to the hydrotreatment stage, said feedstock is added with dimethyl disulfide in order to adjust its sulfur content to 50 ppm by weight.

[0152] [Table 1]

[0153] Table 1: Characteristics of the renewable rapeseed oil feedstock used as feedstock for the hydrotreatment stage.

[0154] Before hydrotreatment of the feedstock, the catalyst is sulfided in situ in the unit, with a distillation gas oil with an additive of 2% by weight of dimethyl disulfide, under a total pressure of 5.1 MPa, a hydrogen / additive gas oil ratio of 700 Nm 3 by m 3 The volume of diesel fuel added per volume of catalyst and per hour is set at 1 h' 1 Sulfurization is carried out for 12 hours at 350°C, with a temperature increase ramp of 10°C per hour.

[0155] After sulfurization, the operating conditions of the unit are adjusted in order to carry out the hydrotreatment of the load:

[0156] - WH (charge volume / catalyst volume / hour): 1 h' 1 , total working pressure: 5.1 MPa,

[0157] - hydrogen / charge ratio: 700 Nm 3 hydrogen / m 3 dump,

[0158] - temperature: 310°C.

[0159] The hydrogen used is supplied by Air Product and has a purity greater than 99.999% by volume.

[0160] Step b): separation of the effluent from step a)

[0161] All of the hydrotreated effluent from step a) is separated using a gas / liquid separator so as to recover a light fraction containing mainly hydrogen, propane, water in the form of vapor, carbon oxides (CO and CO2) and ammonia and a liquid hydrocarbon effluent mainly consisting of linear hydrocarbons. The water present in the liquid hydrocarbon effluent is removed by decantation. The liquid hydrocarbon effluent thus obtained contains an atomic oxygen content of less than 80 ppm by weight, said atomic oxygen content being measured by the infrared adsorption technique described in patent application US2009 / 0018374, and a sulfur content of 2 ppm by weight and a nitrogen content of less than 1 ppm by weight, said nitrogen and sulfur contents being measured respectively by chemiluminescence and UV fluorescence. Said liquid hydrocarbon effluent has a density of 791 kg / m 3The liquid hydrocarbon effluent is composed of paraffins; its composition, measured by gas chromatography, is provided in Table 2.

[0162] [Table 2]

[0163] Table 2: Composition of the liquid hydrocarbon effluent used as feedstock for hydroconversion. Non-compliant example 5: Hydroconversion of the liquid hydrocarbon effluent from example 4 according to a process using hydroconversion catalyst C2 with production of a kerosene cut and a diesel cut without recycling the diesel cut.

[0164] Example 5 is not in accordance with the invention insofar as it illustrates a process not in accordance with the invention but using a catalyst C2 in accordance with the invention in the hydroconversion step c). The hydroconversion catalyst C2 was evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, quality of separation of the effluent) of the industrial process according to the invention. The different steps and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are described below.

[0165] In a reactor regulated in temperature so as to ensure isothermal operation and with a fixed bed loaded with 50 ml of hydroconversion catalyst C2, the catalyst being previously activated by reduction, the hydroconversion of the liquid hydrocarbon effluent from example 4 is carried out. Given the nature of catalyst C2 (noble metal), any injection of sulfur is to be prohibited.

[0166] Catalyst C2 undergoes a reduction step under hydrogen carried out at 5.1 MPa and 400°C (2-hour stage).

[0167] After reduction, the operating conditions of the unit are adjusted in order to carry out the hydroconversion of the liquid hydrocarbon effluent in the following range of operating conditions:

[0168] - WH (charge volume / catalyst volume / hour): 0.5 h' 1 ,

[0169] - total working pressure: 5.1 MPa,

[0170] - hydrogen / charge ratio: 350 Nm 3 hydrogen / m3 dump.

[0171] The hydrogen used and entering the hydroconversion stage is supplied by Air Product, it has a purity greater than 99.999% by volume, it is free of hydrogen sulfide.

[0172] At the outlet of the hydroconversion reactor, the reaction effluent is sent to a gas-liquid separation step carried out using a flash drum operated at a pressure comparable to that of the hydroconversion reactor. The hydrogen-rich gas phase is sent to the gas outlet of the unit. The liquid hydrocarbon phase is pressure-expanded and sent to a stripper so as to stabilize the liquid effluent; the gas phase collected at the top of the stripper is also sent to the gas outlet of the unit.

[0173] At the unit gas outlet, an online analysis by gas chromatography and a gas meter allow the mass of light hydrocarbons produced (essentially hydrocarbons with 1 to 5 carbon atoms) and present in the hydrogen stream to be calculated. The liquid effluent is weighed separately, topped at 120°C to remove the naphtha fraction, the 120°C+ liquid effluent is then reweighed and distilled into two sub-cuts: a 120-X°C kerosene cut and an X°C+ diesel cut, where X is a temperature that can be adapted by the refiner according to its specificities and the targets it wishes to achieve. A classic example of a value of X can be 280 or 290°C.

[0174] Temperature steps in the range 250 to 400°C were carried out in order to determine the maximum yield of kerosene cut 120-X°C. At low temperature, hydroconversion is very low and does not allow to produce a large quantity of kerosene cut 120-X°C. At high temperature, hydroconversion is too strong and leads to the formation of a large quantity of gas and naphtha cut C5-120°C to the detriment of the cut of interest 120-X°C.

[0175] The calculation of the C1-C4 gas yield is determined according to the following calculation:

[0176] C1-C4 yield = [(mass of compounds C1 to C4) / (mass of charge)] x 100, the charge corresponding here to the carbon effluent from example 4.

[0177] The calculation of the yield of C5-120°C naphtha is determined according to the following calculation:

[0178] C5-120°C yield (yield in naphtha cut) = [(mass of C5-120°C liquid effluent) / (mass of feedstock)] x 100, the feedstock here corresponding to the carbon effluent from example 4.

[0179] The calculation of the kerosene yield 120-X°C is determined according to the following calculation:

[0180] Yield 120-X°C (yield in kerosene cut) = [(mass of liquid effluent 120-X°C) / (mass of charge)] x 100, the charge corresponding here to the carbon effluent from example 4.

[0181] The calculation of the diesel yield X°C+ is determined according to the following calculation:

[0182] Yield X°C+ (yield in diesel cut) = [(mass of liquid effluent X°C+) / (mass of charge)] x 100, the charge corresponding here to the carbon effluent from example 4.

[0183] The different masses of gas and liquid effluent correspond to the quantities of gas and liquid accumulated over a certain period of time, typically 24 hours, and the feed mass corresponds to the quantity of feed injected into the hydroconversion reactor during the same period of time.

[0184] A WH 0.5 h -1 , 5.1 MPa total pressure and an H2 / charge ratio of 350 Nm 3 / m 3 , the incremental variation of the temperature between 250 and 400°C made it possible to determine the optimal temperature (noted Tbase) at which the yield of the 120-X°C kerosene cut is maximum.

[0185] At Tbase temperature, for a kerosene cut of 120-280°C, the yield structure is as follows:

[0186] • C1-C4 gas yield: 19% by weight

[0187] • Yield in C5-120°C naphtha cut: 20% by weight

[0188] • Yield in kerosene cut 120-280°C: 43% by weight

[0189] • Yield in diesel cut 280°C+: 18% by weight

[0190] At the same Tbase temperature, for a kerosene cut of 120-290°C, the yield structure is as follows:

[0191] • C1-C4 gas yield: 19% by weight

[0192] • Yield in C5-120°C naphtha cut: 20% by weight

[0193] • Kerosene cutting yield 120-290°C: 47% by weight

[0194] • Yield in diesel cut 290°C+: 14% by weight

[0195] In both cases (X = 280 and 290°C), the characteristics obtained on the kerosene cut of interest and the diesel cut as well as the associated operating conditions are reported in summary table 3. In both cases (X = 280 and 290°C), the kerosene cut 120-X°C has a crystal disappearance point lower than -60°C and a smoke point higher than 25 mm.

[0196] Example 6 compliant: Hydroconversion of the liquid hydrocarbon effluent from Example 4 according to a process using the hydroconversion catalyst C2 with production of a kerosene cut and a diesel cut which is entirely recycled in the isomerization step (c) with a recycle rate of 25% vol.

[0197] Example 6 is in accordance with the invention insofar as it combines a process in accordance with the invention with catalyst C2 in accordance with the invention. The combination of the process and catalyst C2 makes it possible to maximize the yield in the cut of interest, which is kerosene 120-X°C.

[0198] The use of a hydroconversion reactor simulator made it possible to calculate the impact of recycling on cutting yields depending on the recycle rate selected by the refiner.

[0199] The recycle rate is defined as the ratio between the volume of diesel cut X°C+ recycled in the hydroisomerization reactor and the volume of fresh feedstock (from example 4). With a recycle rate of 25% vol, the WH which was 0.5 h -1 in the non-compliant example 5 thus goes to 0.63 h -1 because it integrates 25% vol of X°C+ cut in addition to the fresh feed. Similarly, in order to maintain a constant hydrogen flow rate at the inlet of the hydroisomerization reactor, the H2 / feed ratio increases from 350 Nm 3 / m 3 in the non-compliant example 5 to 278 Nm 3 / m 3 in the case of recycling at 25% vol. The temperature is then adapted to allow, under these operating conditions, to obtain 25% vol of X°C+ cut which will be recycled into the hydroisomerization reactor.

[0200] In the case of a 280°C cut, the temperature required to produce and recycle 25% vol of 280°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 2°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0201] • Total pressure: 5.1 MPa

[0202] • Overall WH: 0.63 h' 1

[0203] • H2 / (fresh charge + recycle 280°C+) ratio: 278 Nm 3 / m 3

[0204] • Temperature: Tbase - 2°C

[0205] Under these conditions, the yields obtained at the terminal of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0206] • C1-C4 gas yield: 10% by weight

[0207] • Yield in C5-120°C naphtha cut: 36% by weight

[0208] • Yield in kerosene cut 120-280°C: 54% by weight

[0209] The operation of recycling the 280°C+ cut with a recycle rate of 25% vol thus made it possible to increase the yield of the kerosene cut of interest 120-280°C from 43% wt (non-compliant example 5) to 54% wt, i.e. a gain of 11% wt in kerosene cut yield. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 3. In the case of a 290°C cut, the temperature required to produce and recycle 25% vol of 290°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 2.8°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0210] • Total pressure: 5.1 MPa

[0211] • Overall WH: 0.63 h' 1

[0212] • H2 / (fresh charge + recycle 290°C+) ratio: 278 Nm 3 / m 3

[0213] • Temperature: Tbase - 2.8°C

[0214] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0215] • C1-C4 gas yield: 8% by weight

[0216] • Yield in C5-120°C naphtha cut: 31% by weight

[0217] • Kerosene cutting yield 120-290°C: 61% by weight

[0218] The operation of recycling the 290°C+ cut with a recycle rate of 25% vol thus made it possible to increase the yield of the kerosene cut of interest 120-290°C from 47% wt (example 5 non-compliant) to 61% wt, i.e. a gain of 14% wt in kerosene cut yield. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 3.

[0219] Example 7 compliant: Hydroconversion of the liquid hydrocarbon effluent from Example 4 according to a process using the hydroconversion catalyst C2 with production of a kerosene cut and a diesel cut which is entirely recycled in the isomerization step (c) with a recycle rate of 50% vol.

[0220] Example 7 is in accordance with the invention insofar as it combines a process in accordance with the invention with catalyst C2 in accordance with the invention. The combination of the process and catalyst C2 makes it possible to maximize the yield in the cut of interest, which is kerosene 120-X°C.

[0221] Example 7 is identical to Example 6 but using a recycle rate of 50% vol instead of 25% vol. In the case of a 280°C cut, the temperature required to produce and recycle 50% vol of 280°C+ cut is equal to the base temperature of the non-compliant Example 5 reduced by 2.5°C. The conditions at the boundaries of the hydroconversion reactor are then as follows:

[0222] • Total pressure: 5.1 MPa

[0223] • Overall WH: 0.75 h' 1

[0224] • H2 / (fresh charge + recycle 280°C+) ratio: 233 Nm 3 / m 3

[0225] • Temperature: Tbase - 2.5°C

[0226] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0227] • C1-C4 gas yield: 9% by weight

[0228] • Yield in C5-120°C naphtha cut: 32% by weight

[0229] • Yield in kerosene cut 120-280°C: 59% by weight

[0230] The operation of recycling the 280°C+ cut with a recycle rate of 50% vol thus made it possible to increase the yield of the kerosene cut of interest 120-280°C from 43% wt (example 5 non-compliant) to 59% wt, i.e. a gain of 16% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-280°C yield is 5% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 3.

[0231] In the case of a 290°C cut, the temperature required to produce and recycle 50% vol of 290°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 3.8°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0232] • Total pressure: 5.1 MPa

[0233] • Overall WH: 0.75 h' 1

[0234] • H2 / (fresh charge + recycle 290°C+) ratio: 233 Nm 3 / m 3

[0235] • Temperature: Tbase - 3.8°C

[0236] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows: C1-C4 gas yield: 7% by weight

[0237] Yield in C5-120°C naphtha cut: 26% by weight

[0238] Kerosene cutting yield 120-290°C: 66% by weight

[0239] The operation of recycling the 290°C+ cut with a recycle rate of 50% vol thus made it possible to increase the yield of the kerosene cut of interest 120-290°C from 47% wt (example 5 non-compliant) to 66% wt, i.e. a gain of 19% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-290°C yield is 5% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 3.

[0240] Example 8 compliant: Hydroconversion of the liquid hydrocarbon effluent from Example 4 according to a process using the hydroconversion catalyst C2 with production of a kerosene cut and a diesel cut which is entirely recycled in the isomerization step (c) with a recycle rate of 100% vol.

[0241] Example 8 is in accordance with the invention insofar as it combines a process in accordance with the invention with catalyst C2 in accordance with the invention. The combination of the process and catalyst C2 makes it possible to maximize the yield in the cut of interest, which is kerosene 120-X°C.

[0242] Example 8 is identical to examples 6 and 7 but taking a recycle rate of 100% vol instead of 25 and 50% vol respectively.

[0243] In the case of a 280°C cut, the temperature required to produce and recycle 100% vol of 280°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 3.3°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0244] • Total pressure: 5.1 MPa

[0245] • Overall WH: 1.0 h' 1

[0246] • H2 / (fresh charge + recycle 280°C+) ratio: 175 Nm 3 / m 3

[0247] • Temperature: Tbase - 3.3°C

[0248] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows: C1-C4 gas yield: 8% by weight

[0249] Yield in C5-120°C naphtha cut: 28% by weight

[0250] Kerosene cutting yield 120-280°C: 64% by weight

[0251] The operation of recycling the 280°C+ cut with a recycle rate of 100% vol thus made it possible to increase the yield of the kerosene cut of interest 120-280°C from 43% wt (example 5 non-compliant) to 64% wt, i.e. a gain of 21% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-280°C yield is 10% wt. Compared to the case with a recycle rate of 50% vol, the gain in 120-280°C yield is 5% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 3.

[0252] In the case of a 290°C cut, the temperature required to produce and recycle 100% vol of 290°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 4.7°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0253] • Total pressure: 5.1 MPa

[0254] • Overall WH: 1.0 h'1

[0255] • H2 / (fresh charge + recycle 290°C+) ratio: 175 Nm 3 / m 3

[0256] • Temperature: Tbase - 4.7°C

[0257] Under these conditions, the yields obtained at the terminals of the HDI hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0258] • C1-C4 gas yield: 6% by weight

[0259] • Yield in C5-120°C naphtha cut: 22% by weight

[0260] • Yield in kerosene cut 120-290°C: 72% by weight

[0261] The operation of recycling the 290°C+ cut with a recycle rate of 100% vol thus made it possible to increase the yield of the kerosene cut of interest 120-290°C from 47% wt (example 5 non-compliant) to 72% wt, i.e. a gain of 25% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-290°C yield is 11% wt. Compared to the case with a recycle rate of 50% vol, the gain in 120-290°C yield is 6% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 3. [Table 3]

[0262] Table 3: Yields and properties of C1-C4 cuts, naphtha, kerosene and diesel for the different examples of the invention using catalyst C2.

[0263] Non-compliant example 9: Hydroconversion of the liquid hydrocarbon effluent from example 4 according to a process using the C3 hydroconversion catalyst with production of a kerosene cut and a diesel cut without recycling the diesel cut.

[0264] Example 9 is not in accordance with the invention insofar as it illustrates a process not in accordance with the invention but using a C3 catalyst in accordance with the invention in hydroconversion step c).

[0265] The C3 hydroconversion catalyst was evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, quality of effluent separation) of the industrial process according to the invention. The different stages and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are described below.

[0266] In a reactor regulated in temperature so as to ensure isothermal operation and with a fixed bed loaded with 50 ml of hydroconversion catalyst C3, the catalyst being previously activated by reduction, the hydroconversion of the liquid hydrocarbon effluent from example 4 is carried out. Given the nature of the catalyst C3 (noble metal), any injection of sulfur is to be prohibited.

[0267] Catalyst C3 undergoes a reduction step under hydrogen carried out at 5.1 MPa and 400°C (2-hour stage).

[0268] After reduction, the operating conditions of the unit are adjusted in order to carry out the hydroconversion of the liquid hydrocarbon effluent in the following range of operating conditions:

[0269] - WH (charge volume / catalyst volume / hour): 0.5 h' 1 ,

[0270] - total working pressure: 5.1 MPa,

[0271] - hydrogen / charge ratio: 350 Nm 3 hydrogen / m3 dump.

[0272] The hydrogen used and entering the hydroconversion stage is supplied by Air Product, it has a purity greater than 99.999% by volume, it is free of hydrogen sulfide. At the outlet of the hydroconversion reactor, the reaction effluent is sent to a gas-liquid separation stage carried out by means of a flash drum operated at a pressure comparable to that of the hydroconversion reactor. The hydrogen-rich gas phase is sent to the gas outlet of the unit. The liquid hydrocarbon phase is pressure-expanded and sent to a stripper so as to stabilize the liquid effluent, the gas phase collected at the top of the stripper is also sent to the gas outlet of the unit.

[0273] At the unit gas outlet, an online analysis by gas chromatography and a gas meter allow the mass of light hydrocarbons produced (essentially hydrocarbons with 1 to 5 carbon atoms) and present in the hydrogen flow to be calculated. The liquid effluent is weighed separately, topped at 120°C to remove the naphtha fraction, the liquid effluent 120°C + is then reweighed and distilled into two sub-cuts: a kerosene cut 120-X°C and a diesel cut X°C+, where X is a temperature that can be adapted by the refiner according to its specificities and the targets they wish to achieve. A classic example of a value of X can be 280 or 290°C.

[0274] Temperature steps in the range 250 to 400°C were carried out in order to determine the maximum yield of kerosene cut 120-X°C. At low temperature, hydroconversion is very low and does not allow to produce a large quantity of kerosene cut 120-X°C. At high temperature, hydroconversion is too strong and leads to the formation of a large quantity of gas and naphtha cut C5-120°C to the detriment of the cut of interest 120-X°C.

[0275] The calculation of the C1-C4 gas yield is determined according to the following calculation:

[0276] C1-C4 yield = [(mass of compounds C1 to C4) / (mass of charge)] x 100, the charge corresponding here to the carbon effluent from example 4.

[0277] The calculation of the yield of C5-120°C naphtha is determined according to the following calculation:

[0278] C5-120°C yield (yield in naphtha cut) = [(mass of C5-120°C liquid effluent) / (mass of feedstock)] x 100, the feedstock here corresponding to the carbon effluent from example 4.

[0279] The calculation of the kerosene yield 120-X°C is determined according to the following calculation: Yield 120-X°C (yield in kerosene cut) = [(mass of liquid effluent 120-X°C) / (mass of charge)] x 100, the charge corresponding here to the carbonaceous effluent from example 4.

[0280] The calculation of the diesel yield X°C+ is determined according to the following calculation:

[0281] Yield X°C+ (yield in diesel cut) = [(mass of liquid effluent X°C+) / (mass of charge)] x 100, the charge corresponding here to the carbon effluent from example 4.

[0282] The different masses of gas and liquid effluent correspond to the quantities of gas and liquid accumulated over a certain period of time, typically 24 hours, and the feed mass corresponds to the quantity of feed injected into the hydroconversion reactor during the same period of time.

[0283] A WH 0.5 h -1 , 5.1 MPa total pressure and an H2 / charge ratio of 350 Nm 3 / m 3 , the incremental variation of the temperature between 250 and 400°C made it possible to determine the optimal temperature at which the efficiency of the 120-X°C kerosene cut is maximum.

[0284] At Tbase temperature - 7°C, for a kerosene cut 120-280°C, the structure-yield is as follows:

[0285] • C1-C4 gas yield: 19% by weight

[0286] • Yield in C5-120°C naphtha cut: 20% by weight

[0287] • Yield in kerosene cut 120-280°C: 44% by weight

[0288] • Yield in diesel cut 280°C+: 17% by weight

[0289] At Tbase temperature - 7°C, for a kerosene cut 120-290°C, the structure-yield is as follows:

[0290] • C1-C4 gas yield: 19% by weight

[0291] • Yield in C5-120°C naphtha cut: 19% by weight

[0292] • Kerosene cutting yield 120-290°C: 47% by weight

[0293] • Yield in diesel cut 290°C+: 15% by weight

[0294] In both cases (X = 280 and 290°C), the characteristics obtained on the kerosene cut of interest and the diesel cut as well as the associated operating conditions are reported in summary table 4. In both cases (X = 280 and 290°C), the kerosene cut 120-X°C has a crystal disappearance point lower than -60°C and a smoke point higher than 25 mm.

[0295] Example 10 compliant: Hydroconversion of the liquid hydrocarbon effluent from Example 4 according to a process using the C3 hydroconversion catalyst with production of a kerosene cut and a diesel cut which is entirely recycled in the isomerization step (c) with a recycle rate of 25% vol.

[0296] Example 10 is in accordance with the invention insofar as it combines a process in accordance with the invention with catalyst C3 in accordance with the invention. The combination of the process and catalyst C3 makes it possible to maximize the yield in the cut of interest, which is kerosene 120-X°C.

[0297] The use of a hydroconversion reactor simulator made it possible to calculate the impact of recycling on cutting yields depending on the recycle rate selected by the refiner.

[0298] The recycle rate is defined as the ratio between the volume of diesel cut X°C+ recycled in the hydroisomerization reactor and the volume of fresh feedstock (from example 4). With a recycle rate of 25% vol, the WH which was 0.5 h -1 in the non-compliant example 9 thus goes to 0.63 h -1 because it integrates 25% vol of X°C+ cut in addition to the fresh feed. Similarly, in order to maintain a constant hydrogen flow rate at the inlet of the hydroisomerization reactor, the H2 / feed ratio increases from 350 Nm 3 / m 3 in the non-compliant example 9 to 278 Nm 3 / m 3 in the case of recycling at 25% vol. The temperature is then adapted to allow, under these operating conditions, to obtain 25% vol of X°C+ cut which will be recycled into the hydroisomerization reactor.

[0299] In the case of a 280°C cut, the temperature required to produce and recycle 25% vol of 280°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 10°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0300] • Total pressure: 5.1 MPa

[0301] • Overall WH: 0.63 h' 1

[0302] • H2 / (fresh charge + recycle 280°C+) ratio: 278 Nm 3 / m 3

[0303] • Temperature: Tbase - 10°C Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0304] C1-C4 gas yield: 10% by weight

[0305] Yield in C5-120°C naphtha cut: 36% by weight

[0306] • Yield in kerosene cut 120-280°C: 54% by weight

[0307] The operation of recycling the 280°C+ cut with a recycle rate of 25% vol thus made it possible to increase the yield of the kerosene cut of interest 120-280°C from 44% wt (example 9 non-compliant) to 54% wt, i.e. a gain of 10% wt in kerosene cut yield. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 4.

[0308] In the case of a 290°C cut, the temperature required to produce and recycle 25% vol of 290°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 9.3°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0309] • Total pressure: 5.1 MPa

[0310] • Overall WH: 0.63 h' 1

[0311] • H2 / (fresh charge + recycle 290°C+) ratio: 278 Nm 3 / m 3

[0312] • Temperature: Tbase - 9.3°C

[0313] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0314] • C1-C4 gas yield: 7% by weight

[0315] • Yield in C5-120°C naphtha cut: 30% by weight

[0316] • Kerosene cutting yield 120-290°C: 63% by weight

[0317] The operation consisting of recycling the 290°C+ cut with a recycle rate of 25% vol thus made it possible to increase the yield of the kerosene cut of interest 120-290°C from 47% wt (example 9 non-compliant) to 63% wt, i.e. a gain of 16% wt in kerosene cut yield. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in table 4. Example 11 compliant: Hydroconversion of the liquid hydrocarbon effluent from example 4 according to a process using the C3 hydroconversion catalyst with production of a kerosene cut and a diesel cut which is entirely recycled in the isomerization step (c) with a recycle rate of 50% vol.

[0318] Example 11 is in accordance with the invention insofar as it combines a process in accordance with the invention with catalyst C3 in accordance with the invention. The combination of the process and catalyst C3 makes it possible to maximize the yield in the cut of interest, which is kerosene 120-X°C.

[0319] Example 11 is identical to example 10 but using a recycle rate of 50% vol instead of 25% vol.

[0320] In the case of a 280°C cut, the temperature required to produce and recycle 50% vol of 280°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 10.5°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0321] • Total pressure: 5.1 MPa

[0322] • Overall WH: 0.75 h' 1

[0323] • H2 / (fresh charge + recycle 280°C+) ratio: 233 Nm 3 / m 3

[0324] • Temperature: Tbase - 10.5°C

[0325] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0326] • C1-C4 gas yield: 10% by weight

[0327] • Yield in C5-120°C naphtha cut: 32% by weight

[0328] • Yield in kerosene cut 120-280°C: 58% by weight

[0329] The operation of recycling the 280°C+ cut with a recycle rate of 50% vol thus made it possible to increase the yield of the kerosene cut of interest 120-280°C from 44% wt (non-compliant example 9) to 58% wt, i.e. a gain of 14% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-280°C yield is 4% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 4. In the case of a 290°C cut, the temperature required to produce and recycle 50% vol of 290°C+ cut is equal to the base temperature of non-compliant example 5 reduced by 10.8°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0330] • Total pressure: 5.1 MPa

[0331] • Overall WH: 0.75 h' 1

[0332] • H2 / (fresh charge + recycle 290°C+) ratio: 233 Nm 3 / m 3

[0333] • Temperature: Tbase - 10.8°C

[0334] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0335] • C1-C4 gas yield: 7% by weight

[0336] • Yield in C5-120°C naphtha cut: 26% by weight

[0337] • Kerosene cutting yield 120-290°C: 67% by weight

[0338] The operation of recycling the 290°C+ cut with a recycle rate of 50% vol thus made it possible to increase the yield of the kerosene cut of interest 120-290°C from 47% wt (example 9 non-compliant) to 67% wt, i.e. a gain of 20% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-290°C yield is 4% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 4.

[0339] Example 12 compliant: Hydroconversion of the liquid hydrocarbon effluent from Example 4 according to a process using the C3 hydroconversion catalyst with production of a kerosene cut and a diesel cut which is entirely recycled in the isomerization step (c) with a recycle rate of 100% vol.

[0340] Example 12 is in accordance with the invention insofar as it combines a process in accordance with the invention with catalyst C3 in accordance with the invention. The combination of the process and catalyst C3 makes it possible to maximize the yield in the cut of interest, which is kerosene 120-X°C.

[0341] Example 12 is identical to Examples 10 and 11 but using a recycle rate of 100% vol instead of 25 and 50% vol respectively. In the case of a 280°C cut, the temperature required to produce and recycle 100% vol of 280°C+ cut is equal to the base temperature of the non-compliant Example 9 reduced by 11.2°C. The conditions at the boundaries of the hydroconversion reactor are then as follows:

[0342] • Total pressure: 5.1 MPa

[0343] • Overall WH: 1.0 h' 1

[0344] • H2 / (fresh charge + recycle 280°C+) ratio: 175 Nm 3 / m 3

[0345] • Temperature: Tbase - 11.2°C

[0346] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows:

[0347] • C1-C4 gas yield: 8% by weight

[0348] • Yield in C5-120°C naphtha cut: 29% by weight

[0349] • Yield in kerosene cut 120-280°C: 63% by weight

[0350] The operation of recycling the 280°C+ cut with a recycle rate of 100% vol thus made it possible to increase the yield of the kerosene cut of interest 120-280°C from 44% wt (example 9 non-compliant) to 63% wt, i.e. a gain of 19% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-280°C yield is 9% wt. Compared to the case with a recycle rate of 50% vol, the gain in 120-280°C yield is 5% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 4.

[0351] In the case of a 290°C cut, the temperature required to produce and recycle 100% vol of 290°C+ cut is equal to the base temperature of non-compliant example 9 reduced by 11.6°C. The conditions at the terminals of the hydroconversion reactor are then as follows:

[0352] • Total pressure: 5.1 MPa

[0353] • Overall WH: 1.0 h' 1

[0354] • H2 / (fresh charge + recycle 290°C+) ratio: 175 Nm 3 / m 3

[0355] • Temperature: Tbase - 11.6°C

[0356] Under these conditions, the yields obtained at the terminals of the hydroconversion section (hydroconversion reactor + distillation + recycling) are as follows: C1-C4 gas yield: 7% by weight

[0357] Yield in C5-120°C naphtha cut: 23% by weight

[0358] Kerosene cutting yield 120-290°C: 70% by weight

[0359] The operation of recycling the 290°C+ cut with a recycle rate of 100% vol thus made it possible to increase the yield of the kerosene cut of interest 120-290°C from 47% wt (example 9 non-compliant) to 70% wt, i.e. a gain of 23% wt in kerosene cut yield. Compared to the case with a recycle rate of 25% vol, the gain in 120-290°C yield is 7% wt. Compared to the case with a recycle rate of 50% vol, the gain in 120-290°C yield is 3% wt. The properties of the kerosene cut are only very slightly impacted by the introduction of this recycle as shown in Table 4.

[0360] [Table 4]

[0361] Table 4: Yields and properties of C1-C4 cuts, naphtha, kerosene and diesel for the different examples of the invention using the C3 catalyst.

Claims

CLAIMS 1. A process for treating a feedstock from a renewable source to produce a kerosene cut, said process comprising at least the following steps, and preferably consisting of the following steps: a) a step of hydrotreating said feedstock in the presence of a fixed-bed catalyst, said catalyst comprising a hydrogenating function and an oxide support, at a temperature of between 200 and 450°C, at a pressure of between 1 MPa and 10 MPa, at an hourly space velocity of between 0.1 h-1 and 10 h-1 and in the presence of a total quantity of hydrogen mixed with the feedstock such that the hydrogen / feedstock ratio is between 70 and 2000 Nm3 of hydrogen / m3 of feedstock, b) a step of separating at least a portion of the effluent from step a) into at least one light gaseous fraction, at least one hydrocarbon liquid effluent, and at least one aqueous liquid effluent,c) a step of hydroconversion of at least a portion of the liquid hydrocarbon effluent from step b) in the presence of at least one fixed-bed bifunctional hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one noble metal from group VIII of the periodic table chosen from platinum and palladium and a support comprising at least one zeolite chosen from zeolites with structural code MTW, and zeolite IZM-2 alone or as a mixture, and at least one binder, said hydroconversion step being carried out at a temperature of between 250°C and 500°C, at a pressure of between 1 MPa and 10 MPa, at an hourly space velocity of between 0.1 and 10 h-1 and in the presence of a total quantity of hydrogen mixed with the feedstock such that the hydrogen / feedstock ratio is between 70 and 1000 Nm3 / m3 of feedstock,d) a step of separating at least a portion of the effluent from step c) which makes it possible to separate at least a gaseous fraction and at least one liquid hydrocarbon effluent, e) a step of fractionating the hydrocarbon effluent from step d) into at least one kerosene fraction, at least one heavy fraction having an initial boiling point of between 250 and 300°C, Said process comprising a step of recycling all or part of said heavy cut having an initial boiling point of between 250 and 300°C from said fractionation step e) in said hydroconversion step c).

2. Method according to claim 1 in which the feedstock from renewable sources is chosen from oils and fats of vegetable or animal origin, used cooking oils, oils of microbial origin, fish oils, long paraffins from the Fischer-Tropsch process, raw or having undergone prior treatment, or mixtures of such feedstocks, containing triglycerides and / or free fatty acids and / or esters.

3. Method according to one of claims 1 or 2 wherein in step a), the feed is brought into contact with a fixed bed catalyst at a temperature of between 220 and 350°C, at a pressure of between 1 MPa and 6 MPa, at an hourly space velocity of between 0.1 h -1 and 10 a.m. -1 The charge is brought into contact with the catalyst in the presence of hydrogen and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 150 and 1000 Nm3 of hydrogen / m 3 dump.

4. Process according to one of claims 1 to 3 in which the hydrotreatment catalyst used in step a) comprises a hydro-dehydrogenating function comprising at least one metal from group VIII and / or group VIB, taken alone or as a mixture and a support chosen from the group formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals.

5. Process according to one of claims 1 to 4 in which the hydroconversion step c) operates at a temperature of between 250°C and 450°C, and very preferably, between 250 and 400°C, at a pressure of between 2 MPa and 10 MPa and very preferably, between 3 MPa and 9 MPa, at an hourly volumetric flow rate advantageously of between 0.2 and 7 h -1 and very preferably, between 0.5 and 5 h' 1, at a hydrogen flow rate such that the hydrogen / charge volume ratio is advantageously between 100 and 1000 normal m 3 of hydrogen per m 3 load and preferably between 150 and 1000 normal m 3 of hydrogen per m 3 dump.

6. Process according to one of claims 1 to 5 in which the group VIII metal of the catalyst used in step c) is platinum.

7. Process according to one of claims 1 to 6 in which the zeolite with structural code MTW used in the catalyst of step c) is ZSM-12 8. Process according to one of claims 1 to 7 in which the catalyst used in step c) comprises a hydrogenating phase comprising platinum and a support comprising an IZM-2 zeolite alone and an alumina binder.

9. Process according to one of claims 1 to 7 in which the catalyst used in step c) comprises a hydrogenating phase comprising platinum and a support comprising a ZSM-12 zeolite alone and an alumina binder.

10. Process according to one of claims 1 to 9 in which the fractionation step e) is carried out in a distillation column or in a steam stripping step followed by a distillation column.

11. Process according to one of claims 1 to 10 in which the heavy cut having an initial boiling point of between 250 and 300°C from fractionation step e) is mixed with all or part of the liquid hydrocarbon effluent from step b) upstream of the hydroconversion reactor and then this mixture is injected into the first catalytic bed of step c), in the case where said hydroconversion step comprises several catalytic beds.

12. Process according to one of claims 1 to 10 in which the heavy cut having an initial boiling point of between 250 and 300°C from fractionation step e) is injected into an intermediate bed of the hydroconversion reactor in the case where said hydroconversion step comprises several catalytic beds.

13. Process according to one of claims 1 to 12 in which the process does not include a hydrocracking step downstream of step c) of hydroconversion of part or all of the effluent from step c).

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