PROCESS FOR THE PRODUCTION OF RENEWABLE KEROSENE COMPRISING TWO SUCCESSIVE HYDROCONVERSION STAGES AND USING A SPECIFIC CATALYST FOR EACH STAGE.

Abstract

The present invention describes a process for treating a feedstock from a renewable source for the production of a kerosene cut or a diesel base, employing a chain of two specific bifunctional catalysts, comprising at least one metal from Group VIII and / or Group VIB, taken alone or in mixture, and an acidic support, preferably silica alumina, for the first catalyst, and at least one metal from Group VIII and / or Group VIB and a support comprising one or more specific zeolites selected from structural type FER, EUO, *MRE, MTW, MOR and IZM-2 zeolite, alone or in mixture, and preferably selected from structural code MTW zeolites and IZM-2 for the second catalyst, in two successive hydroconversion steps without a final fractionation step, allowing a kerosene cut to be obtained to the specifications of ASTM D7566 with a maximized yield in said cut.said process also allows for the alternative production of a diesel base that can be incorporated into the diesel pool, without loss of yield compared to a process not implementing the specific catalytic sequence according to the invention. Figure to be published: 1,

Description

Title of the invention: METHOD FOR PRODUCING RENEWABLE KEROSENE COMPRISING TWO STAGES SUCCESSIVE HYDROCONVERSION AND USING A SPECIFIC CATALYST FOR EACH OF THE STAGES. Scope of the invention

[0001] The search for new renewable energy sources for fuel production is a major challenge in order to both meet fuel demand and take into account environmental concerns.

[0002] As such, the use of feedstocks from renewable sources in fuels has seen a significant resurgence of interest in recent years. Examples of these feedstocks include vegetable oils (e.g., palm, rapeseed, soybean), animal fats, pine oil, used cooking oil, microbial oils (e.g., from algae), fish oil, and long paraffins (waxes) from the Fischer-Tropsch process, either crude or pre-treated, as well as mixtures of such feedstocks. Most of these feedstocks contain chemical structures such as triglycerides, esters, or fatty acids, the structure and hydrocarbon chain length of which are compatible with the hydrocarbons present in diesel fuel and kerosene.

[0003] One possible route is the catalytic transformation of the feedstock from a renewable source into deoxygenated paraffinic fuel in the presence of hydrogen (hydrotreating). Many metallic or sulfide catalysts are known to be active for this type of reaction.

[0004] These hydrotreating processes for 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.

[0005] The use of solids based on transition metal sulfides allows the production of paraffins from ester-type molecules via two reaction pathways:

[0006] - hydrodeoxygenation leading to the formation of water by consumption of hydrogen and the formation of hydrocarbons with a carbon number (Cn) equal to that of the initial fatty acid chains,

[0007] - decarboxylation / decarbonylation leading to the formation of carbon oxides (carbon monoxide and carbon dioxide: CO and CO2) and the formation of hydrocarbons with one less carbon (Cn.i) compared to the initial fatty acid chains.

[0008] The liquid effluent from these hydrotreating processes, after separation, consists essentially of n-paraffins and is substantially free of sulfur, nitrogen, and oxygen impurities. After hydrotreating and gas separation, the sulfur content is typically between 1 and 20 ppm wt. The nitrogen content is generally between 0.2 and 30 ppm wt., and the oxygen content is generally less than 2000 ppm wt. The paraffins typically have a carbon atom number between 9 and 25, which is mainly dependent on the composition of the feedstock to be hydrotreated.

[0009] However, this liquid effluent cannot generally be incorporated as is into the kerosene or diesel pool, particularly due to insufficient cold-weather properties and / or excessively high boiling points. Indeed, the paraffins present lead to high pour points and therefore to freezing phenomena for low-temperature applications. For example, eicosane (linear paraffin with 20 carbon atoms, C2OH) has a boiling point of 340°C and a melting point of 37°C. The boiling point of eicosane is thus compatible with incorporation into a diesel pool, but its melting point can generate freezing problems and limit its use. For example, the maximum filter plugging point for winter diesel is -15°C in France according to standard NF EN 590.Furthermore, the boiling point of eicosane makes it unsuitable for incorporation into the kerosene pool, for which the final temperature of the D86 distillation curve must be less than 300°C according to ASTM D1655.

[0010] Depending on the incorporation rate and the preferred fuel pool (diesel or kerosene) targeted, a hydroconversion step (hydroisomerization and / or hydrocracking reactions) may be necessary to transform the linear paraffins in the hydrotreated liquid effluent. Hydroisomerization converts a linear paraffin into a branched paraffin while conserving the number of carbon atoms in the molecule. This improves the cold properties of the effluent because branched paraffins exhibit better cold properties than linear paraffins. For example, nonadecane has a melting point of 32°C, while one of its monobranch isomers, 7-methyloctadecane, 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, the hydrocracking of one molecule of eicosane can lead to the production of two molecules of 2-methylnonane. The boiling point of 2-methylnonane is 167°C, which is compatible with its incorporation into the kerosene pool. The hydroconversion step is carried out on a bifunctional catalyst exhibiting both a hydro / dehydrogenating function and a Brønsted acid function. The conditions... Operating procedures 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.

[0011] The appropriate choice of acid phase promotes the isomerization of long linear paraffins and minimizes cracking. Thus, the shape selectivity of one-dimensional medium-pore (10 MR) zeolites such as ZSM-22, ZSM-23, NU-10, ZSM-48, and ZBM-30 makes their use particularly suitable for obtaining catalysts selective for isomerization. Other acidic phases, both zeolitic and non-zeolitic, such as halogenated aluminas (particularly chlorinated or fluorinated), phosphorus aluminas, silica-aluminas, or silicified aluminas, can also be used.

[0012] However, it is well known that factors other than the acid phase have an impact on the activity and selectivity of a bifunctional catalyst. Hydroisomerization and hydrocracking of normal paraffins have thus 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 the n-paraffin being dehydrogenated to n-olefin in the hydro-dehydrogenating phase and then, after diffusion to the acid phase, being protonated to carbenium ions. After structural rearrangement and / or [3-cleavage], the carbenium ions desorb from the acid phase as olefins after deprotonation. Then, after diffusion to the hydro-dehydrogenating phase, the olefins are hydrogenated to form the final reaction products.It is therefore necessary to have a sufficiently active hydro / dehydrogenating function with respect to the acid function in order to, on the one hand, rapidly supply the acid phase with olefins and, on the other hand, rapidly hydrogenate the olefinic intermediates after their reaction with the acid phase. This allows, on the one hand, for maximizing the activity of the catalyst and, on the other hand, for favoring hydroisomerization over hydrocracking when the former reaction is desired, or for limiting the production of excessively light cracking products 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.

[0013] The proximity between the two functions of the catalyst 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 in which platinum is deposited on alumina is consistently more selective for isomerization than the catalyst in which platinum is deposited within the zeolite. Based on these results, those skilled in the art are therefore inclined to favor localizing the hydrogenating function on the alumina matrix rather than on the acid phase to improve isomerization selectivity. From an activity standpoint, the localization of platinum on the alumina matrix has a variable impact depending on the long paraffin wax considered: a positive impact for n-decane, a marginal impact for n-nonadecane, and a negative impact for pristane.

[0014] Noble metals (Pt, Pd) or VIB transition metals (Mo, W) combined with group VIII transition metals (Ni, Co) can act as hydrogenating agents for the catalyst. The noble metals are used in their reduced form, while the transition metals are used in their sulfide form. For the latter, there is a known synergistic effect between VIB and group VIII transition metals, generally attributed to the decoration of the VIB sulfide phases by the group VIII transition metals. These are referred to as nickel- or cobalt-promoted molybdenum or tungsten sulfide phases ("CoMoS", "NiMoS", "NiWS"). This synergistic effect results in an increase in the catalytic activity of the promoted phase compared to an unpromoted phase.

[0015] The choice of the nature of the hydrogenating function, whether noble metal or sulfide, depends on various criteria, both economic (the price of noble metals is significantly higher than that of transition metals in groups VIB and VIII) and chemical (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 significant (C. Marcilly, Catalyse acido-basique, volume 2, 2003, Technip editions).

[0016] US patent application 2022 / 0127537 discloses a hydrotreating process for a renewable feedstock. The process comprises a hydrotreating step of the feedstock in the presence of hydrogen and a hydrotreating catalyst to deoxygenate the feedstock and thus produce a hydrotreated effluent. The process also comprises a hydroisomerization step, in the presence of hydrogen and a hydroisomerization catalyst, of an effluent derived from the hydrotreated effluent to obtain a hydroisomerized effluent. In the case of a two-step process, the catalyst The hydroisomerization agent used may include 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, silica, amorphous 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, offretite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.

[0017] In a first embodiment, the process comprises a step of separating the hydrotreated and / or isomerized effluent to obtain a steam effluent and a liquid effluent; a distillation step of the liquid effluent or the hydroisomerized effluent produces a kerosene fraction and a diesel fraction. The diesel fraction undergoes a further hydrocracking step to obtain a hydrocracking effluent comprising a kerosene fraction. The hydrocracking catalyst comprises a Group VIII metal or a Group VIB metal on an amorphous silica-alumina support or on a zeolite support. In the case where the hydrocracking catalyst is zeolite, the active phase comprises a Group VIII metal, optionally in combination with a Group VIB metal.Zeolites are preferably chosen from mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite and faujasite, and preferably zeolites B, X, Y and L, and preferably zeolite Y. The said process thus allows the co-production of diesel and kerosene fractions which are separated by a distillation step and the resulting diesel fraction then undergoes a hydrocracking step.

[0018] In a second embodiment, the process comprises obtaining a diesel fraction from the hydroisomerized effluent. This diesel fraction is characterized in that the concentration of normal paraffin at a given carbon number is at least twice the corresponding concentration of normal paraffin in the hydroisomerized effluent. The diesel fraction undergoes a hydrocracking or hydroisomerization step to obtain a hydrocraced effluent comprising a kerosene fraction. In contrast to the present invention, the process therefore comprises either a hydrocracking step of the diesel fraction obtained or a recycling step of the diesel fraction in the hydroisomerization step.

[0019] In a third embodiment, 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 fraction and a diesel fraction; and a step of hydrocracking said diesel fraction to obtain a hydrocraced effluent comprising a fraction kerosene. This process therefore allows the co-production of diesel and kerosene fractions which are separated by a distillation step and the resulting diesel fraction then undergoes a hydrocracking step.

[0020] Patent application WO23126564 describes a process for producing kerosene comprising a pretreatment step of the renewable feedstock to reduce impurities, a hydrodeoxygenation step of the pretreated feedstock, a gas / liquid separation step of the hydrotreated effluent and then hydroisomerization of the separated liquid effluent in the presence of a hydroisomerization catalyst comprising a noble metal (preferably platinum) and a support comprising a specific 12MR zeolite (preferably ZSM-12) having a pore size of less than 0.7 nm, an acidity between 180 micromol / g and 500 micromol / g measured by the NH3-TPD method in which the acidity is calculated from the amount of NH3 adsorbed at 200 °C and desorbed between 100 °C and 500 °C.The process also optionally includes a step for stabilizing the hydroisomerized effluent and a step for separating the hydroisomerized or stabilized effluent to recover a C10-C16 kerosene cut and a C5-C9 gasoline cut. Patent application WO23126564 does not describe the use of a specific catalyst sequence in the hydroisomerization step.

[0021] Application WO15063213 describes a process for converting a specific paraffinic feed, characteristic of a paraffinic feed from the Fischer Tropsch process, comprising at least 50% of 370°C+ compounds having a paraffin content of at least 60%, an aromatics content of less than 1%, a naphthene content of less than 2%, a nitrogen content of less than 0.1%, and a sulfur content of less than 0.1%, wherein the feed is converted to obtain an effluent at least partially isomerized by a series of two catalysts, a first catalyst 1 preferably comprising a Group VIII metal supported on an amorphous acidic support, preferably Pt / SiAl, and a second catalyst 2 that is more isomerizing and less cracking than the first catalyst 1, preferably comprising a Group VIII metal and a molecular sieve selected from MTW, MTT, TON, or zeolites. ZSM-48 and preferably MTW type zeolites.The preferred catalyst 2 comprises Pt or Pd, MTW zeolite, and a silica binder. A highly preferred catalyst 2 is an ammonium hexafluorosilicate-treated Pt / ZSM12 catalyst comprising a silica binder.

[0022] The process then includes a step of separating the isomerized effluent into a diesel fraction, a kerosene fraction and a residual fraction boiling over the middle distillates.

[0023] Patent application WO10077476 describes a process for producing diesel fuel from vegetable oil comprising a step of hydrodeoxygenating the feedstock for The process involves producing a paraffinic effluent, a hydroisomerization and selective hydrocracking step of a portion of said paraffinic effluent, followed by a final separation step of a naphtha cut, an LPG cut, and a diesel cut comprising paraffins, and a recycled effluent comprising paraffins and boiling in the diesel range, wherein the diesel cut and the recycled effluent do not have the same composition, and optionally a kerosene cut. The recycled effluent (boiling in the diesel range) is then recycled back into the hydroisomerization and selective hydrocracking step. Patent application WO10077476 does not describe the use of a specific catalyst sequence in the hydroisomerization step.Furthermore, unlike the present invention, the described process includes a final fractionation step enabling the production of a diesel cut and a kerosene cut, and the recycling of a portion of the diesel cut in the hydroisomerization step.

[0024] Patent application WO09120242 describes a process for producing hydrocarbons from a renewable feedstock by hydrotreating and deoxygenating said feedstock to produce a paraffinic effluent, isomerizing part of the paraffinic effluent in the presence of an isomerization catalyst and then selectively cracking in the presence of a selective hydrocracking catalyst to recover the hydrocarbon effluent, the selective cracking can be done before, after or simultaneously with the isomerization and the isomerization and hydrocracking reactions can be carried out at the same time in the presence of one and the same catalyst or separately in the presence of 2 different catalysts.

[0025] The described isomerization catalysts 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, silica, amorphous 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, offretite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.

[0026] The hydrocracking catalysts described comprise a silica-alumina-based support, a Beta zeolite and alumina or silica-alumina, or comprise a Y zeolite, or a zeolite selected from ZSM-5, ZSM-23, ZSM-11 and ZSM-22 zeolites and ferrierite. The metallic element is selected from platinum, palladium, nickel, iridium, rhodium, or a mixture thereof.

[0027] The process also provides for the recycling of a fraction of the n-paraffins obtained after hydrotreatment in the hydrotreatment step with a volume recycling rate of the recycle on the feed of between 2 and 8 vol.

[0028] Patent CN116024003B describes a process for producing kerosene in a process comprising a hydrotreating step, a gas / liquid separation step of the hydrotreated effluent, followed by hydrocracking of the separated liquid effluent in the presence of a hydrocracking catalyst comprising a Group VIII metal with an oxide content of between 0.1 and 5 wt% and a support comprising alumina and silica alumina. The conversion level in the hydrocracking unit, operating at over 270°C, is controlled to be in the range of between 50% and 90%. The process also includes a hydroisomerization step of the hydrocracking effluent in the presence of a hydroisomerization catalyst comprising a metal selected from Co, Ni, Pd, Pt, Mo, and W with a corresponding oxide content of between 0.2 and 5 wt% and a support comprising alumina and a molecular sieve. mesoporous selected from the zeolites ZSM-22, Nu-10, theta-1, ISL1, ZSM-23, SAPO-11, SAPO-31, SAPO-41,The molecular sieve content is between 20 and 80% by weight relative to the catalyst mass, and the alumina content is between 15 and 75% by weight. The hydroisomerized effluent then undergoes a gas / liquid separation step, and the liquid is fractionated to obtain a naphtha cut and a kerosene cut.

[0029] On the contrary, the present invention provides for the implementation of a very specific chain of catalysts different from that described in CN116024003B.

[0030] In the sense of the present invention, the different embodiments presented can be used alone or in combination with each other, without limitation of combination.

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

[0032] In the following text, the term naphtha cut or naphtha refers to the hydrocarbon fraction having a boiling point lower than that of the middle distillates cut. The middle distillates cut generally has an initial cutting point between 120 and 180°C, preferably 120°C. The naphtha cut can have boiling points ranging from that of hydrocarbon compounds having 5 carbon atoms per molecule (or a boiling point of 36°C) up to 216°C and includes the gasoline cut.

[0033] Throughout the rest of the text, the term kerosene cup or kerosene means the cup having initial and final boiling points between 120 and 300°C and means by diesel cut or diesel, a cut having initial and final boiling points within a range of 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 IUP AC classification, and group VIB to the metals in column 6.

[0035] In the following 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 clarification will be provided by the present invention.

[0036] In the present 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 fraction, the applicant discovered that the use of a chain of two specific bifunctional catalysts, comprising a hydrogenating / dehydrogenating phase including at least one metal from Group VIII and / or Group VIB of the periodic table, taken alone or in mixtures, and an acidic support, preferably silica-alumina, preferably a silica-alumina for the first catalyst, and at least one metal from Group VIII and / or Group VIB and a support including one or more specific zeolites selected from structurally type FER, EUO, *MRE, MTW, MOR, and IZM-2 zeolites, taken alone or in mixtures, and preferably selected from structurally type MTW zeolites and IZM-2 for the second catalyst, implemented in two successive hydroconversion steps without a fractionation step final,was of great interest. In particular, the process according to the invention makes it possible to obtain a kerosene cut meeting the specifications of ASTM D7566 with maximized yield in said cut. Furthermore, said process also makes it possible to obtain, alternatively, a diesel base suitable for incorporation into the diesel pool, without loss of yield compared to a process not implementing the specific catalytic sequence claimed. Object of the invention,

[0039] More specifically, the present invention relates to a process for treating a feedstock from a renewable source to produce alternatively a diesel base or a kerosene cut, comprising at least the following steps and preferably consisting of:

[0040] a) a hydrotreating step of said feed in the presence of a fixed-bed catalyst, said catalyst comprising a hydrogenating function and an oxide support, at a temperature between 200 and 450°C, at a pressure between 1 and 10 MPa, at a spatial velocity between 0.1 and 10 h 1 and in the presence of a total quantity of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 2000 Nm3 of hydrogen / m3 of feed,

[0041] b) a step of separating at least a portion of the effluent from step a) into at least a light gaseous fraction, at least a hydrocarbon liquid effluent consisting of n-paraffins, and at least an aqueous liquid effluent,

[0042] c) a first hydroconversion step of at least a portion of the hydrocarbon liquid effluent from step b) in the presence of a first bifunctional fixed-bed hydroconversion catalyst, said catalyst comprising at least one metal from Group VIII and / or Group VIB of the periodic table, taken alone or in mixture, and an acidic solid support comprising at least one silica alumina and / or one or more zeolites, said first hydroconversion step being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at a spatial velocity between 0.1 and 10 h⁻¹, and in the presence of a total quantity of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm³ / m³ of feed,

[0043] d) a second hydroconversion step of all the hydrocarbon effluent from step c) in the presence of a second fixed-bed hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one metal from Group VIII and / or Group VIB of the periodic table and a support comprising at least one zeolite selected from the structural type zeolites FER, EUO, *MRE, MTW, MOR, and IZM-2 zeolite alone or in mixture, and at least one binder, said second hydroconversion step being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at a spatial velocity between 0.1 and 10 h⁻¹ and in the presence of a total quantity of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm³ / m³ of feed,

[0044] e) a separation step of at least a portion of the effluent from step c) which allows the separation of at least a gaseous fraction, and at least a hydrocarbon liquid effluent,

[0045] f) a stabilization step of at least a portion of the hydrocarbon liquid effluent from step d), so as to separate three hydrocarbon fractions: a light gaseous fraction, a naphtha hydrocarbon fraction and alternatively a diesel base or a kerosene fraction,

[0046] no recycling step in steps c) and / or d) of hydroconversion, of all or part of the hydrocarbon liquid effluents from step d) of hydroconversion, of the effluents from step e) of separation and / or of the effluents from step f) of stabilization being implemented.

[0047] An advantage of the present invention is to provide a process for treating a feedstock from a renewable source to produce a kerosene cut conforming to the specifications of ASTM D7566, directly at the outlet of the second hydroconversion stage, without needing to fractionate the effluent obtained after separation of the naphtha cut, while maximizing the yield of said kerosene cut, through the implementation of two successive hydroconversion stages of the hydrocarbon effluent from the hydrotreatment stage and by the use of a specific catalyst for each hydroconversion stage.

[0048] Another advantage of the present invention is to allow the processing of heavier renewable source feedstocks within said process through the use of a specific catalyst promoting the cracking of long paraffins in the first hydroconversion step followed by a second, more selective hydroconversion step in hydroisomerization through the use of a second specific catalyst.

[0049] Another advantage of the present invention is to provide a process for treating a feedstock from a flexible renewable source, in which, depending on the preferred fuel pool (diesel or kerosene), the operating conditions implemented in the hydroconversion step can advantageously be adjusted to produce either a diesel or biodiesel blend, or a kerosene or bio-kerosene blend. Detailed description of the invention Feedstocks

[0050] The present invention is particularly dedicated to the preparation of diesel and / or kerosene fuel bases corresponding to the new environmental standards, from feedstocks from renewable sources.

[0051] The feedstocks from renewable sources used in the process according to the present invention are advantageously selected from vegetable oils (e.g., palm, rapeseed, soybean), animal fats, used cooking oils, oils of microbial origin (e.g., from algae), fish oils, pine oils, long paraffins (waxes) from the Fischer-Tropsch process, crude or pre-treated, or mixtures of such feedstocks, containing triglycerides and / or free fatty acids and / or esters. The vegetable oils may advantageously be crude or refined, wholly or partially, and derived from the following plants: rapeseed, sunflower, soybean, palm, palm kernel, olive, coconut, jatropha, this list not being exhaustive. Algae 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 industries.

[0052] These fillers essentially contain triglyceride-type chemical structures, also known to those skilled in the art as fatty acid triesters, as well as free fatty acids. A fatty acid triester is thus composed of three fatty acid chains. These fatty acid chains, whether in triester or free fatty acid form, 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 5 to 6 unsaturations per chain.

[0053] The molecules present in the feeds 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 feeds, the degree of unsaturation, expressed as the number of unsaturations per hydrocarbon fatty chain, is advantageously between 0 and 6.

[0054] Fillers from renewable sources also contain minor constituents, generally representing 1 to 5% of their composition. These minor constituents may be of natural origin or alteration and / or contaminant compounds characterized by various types of chemical structures (aldehydes, alcohols, carboxylic acids, hydrocarbons, etc.), different polarities, and highly variable molecular weights. Among the natural constituents, so-called unsaponifiable compounds may represent 0.1 to 3% of the filler and are essentially composed of sterols, tocopherols, tocotrienols, triterpenes (e.g., squalene), natural pigments (e.g., carotenoids and chlorophylls), fatty alcohols, waxes, and hydrocarbons, among others, with molecular weights up to C40. Alteration compounds and contaminants originate from the extraction, refining, and utilization processes of the fillers.Examples of alteration compounds include the products of oxidation and thermal degradation of lipids such as free fatty acids, peroxides, fatty acid dimers, and triglyceride polymers.

[0055] Feedstocks from renewable sources generally also contain various impurities, including heteroatoms such as nitrogen. Nitrogen content in vegetable oils is generally between 1 and approximately 1000 ppm by weight, depending on their nature. Process and catalysts Step a)

[0056] Advantageously, the feedstock may undergo a pretreatment or pre-refining step prior to step a) of the process according to the invention in order to remove, by appropriate treatment, contaminants such as metals, alkali compounds (for example, on ion-exchange resins), alkaline earth metals, and phosphorus. Appropriate treatments may, for example, be thermal and / or chemical treatments well known to those skilled in the art.

[0057] According to step a) of the process according to the invention, the feed, optionally pretreated, is brought into contact with a fixed-bed catalyst at a temperature between 200 and 450°C, preferably between 220 and 350°C, most preferably between 220 and 320°C, and even more preferably between 220 and 310°C. The pressure is between 1 and 10 MPa, most preferably between 1 and 6 MPa, and even more preferably between 1 and 4 MPa. The hourly spatial velocity, i.e., the volume of feed per volume of catalyst per hour, is between 0.1 and 10 h⁻¹. The feed is brought into contact with the catalyst in the presence of hydrogen. The total amount of hydrogen mixed with the feed is such that the hydrogen / feed ratio is between 70 and 2000 Nm3 of hydrogen / m3 of feed and preferably between 150 and 1000 Nm3 of hydrogen / m3 of feed.

[0058] In step a) of the process according to the invention, the fixed-bed catalyst is advantageously a hydrotreating catalyst comprising a hydro-dehydrogenating function including at least one metal from Group VIII and / or Group VIB, alone or in mixtures, and a support selected 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 selected from the group formed by boron oxide, zirconia, titanium oxide, and phosphoric anhydride. The preferred support is an alumina support, and most preferably a q, ô, or y alumina support.

[0059] Said catalyst is advantageously a catalyst comprising group VIII metals preferably selected from nickel and cobalt, taken alone or in mixture, preferably in association with at least one group VIB metal preferably selected from molybdenum and tungsten, taken alone or in mixture.

[0060] The content of metal oxides of groups VIII and preferably of 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 VIB and preferably of molybdenum trioxide is advantageously between 1% and 30% by weight of molybdenum trioxide (MoO3), preferably from 5% to 25% by weight, the percentages being expressed as % weight relative to the total mass of the catalyst.

[0061] The total content of metal oxides of groups VIB 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.

[0062] The catalyst used in step a) of the process according to the invention must advantageously be characterized by a high hydrogenating capacity so as to direct the selectivity of the reaction as much as possible towards hydrogenation that conserves the number of carbon atoms in the fatty acid chains, i.e., the hydrodeoxygenation pathway, in order to maximize the yield of hydrocarbons suitable for the distillation of kerosene and / or diesel fuels. Therefore, it is preferably carried out at a relatively low temperature. Maximizing the hydrogenating function also makes it possible to limit polymerization and / or condensation reactions leading to the formation of coke, which would degrade the stability of the catalytic performance. Preferably, a Ni or NiMo type catalyst is used.

[0063] Group VIII and / or Group VIB metals are introduced by any method known to those skilled in the art, for example, by dry impregnation of the substrate with the metal precursor(s) dissolved in a solvent, which may be water. One or more organic compounds may also be added during this impregnation step or in a subsequent impregnation step. These organic compounds may contain oxygen and / or nitrogen and / or sulfur. For example, in the case of using oxygenated compounds, examples include compounds comprising one or more functional groups selected from among a carboxyl group, alcohol, ether, ketone, ester, or carbonate, or even furanic compounds or sugars.

[0064] The catalyst used in step a) of the hydrotreating process according to the invention may also advantageously contain a dopant element selected from phosphorus and boron, alone or in a mixture. This dopant element may be introduced into the matrix or, preferably, deposited on the support. Silicon may also be deposited on the support, alone or with phosphorus and / or boron and / or fluorine.

[0065] The weight content of the oxide of said dopant element is advantageously less than 20% by weight and preferably less than 10% by weight and is advantageously at least 0.001% by weight.

[0066] Preferred catalysts are the catalysts described in patent application FR 2 943 071 describing catalysts having high selectivity for hydrodeoxygenation reactions.

[0067] Other preferred catalysts are the catalysts described in patent application EP 2 210 663 describing supported or bulk catalysts comprising a phase active consisting of a sulfur compound from group VIB, the element of group VIB being molybdenum.

[0068] The metals of the catalysts used in step a) of hydrotreating of the process according to the invention are sulfide metals or metallic phases and preferably sulfide metals.

[0069] It would not depart from 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 downward flow of liquid.

[0070] Said step a) of hydrotreatment allows hydrogenation, hydrodeoxygenation, hydrodeazotation and hydrodesulfurization of said feed. Step b)

[0071] In accordance with step b) of the process according to the invention, a separation step is carried out to separate at least a portion, and preferably all, of the effluent from step a). This step b) separates at least one hydrogen-rich, so-called light gaseous fraction, at least one hydrocarbon liquid effluent consisting of n-paraffins, and at least one aqueous liquid effluent.

[0072] 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, i.e., the C1 to C4 compounds 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, 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 w / w and a nitrogen content of less than 2 ppm w / w.

[0073] The hydrocarbon liquid effluent consisting of n-paraffins from said step b) advantageously comprises a content of compounds boiling at a temperature above 370°C strictly greater than 50%, preferably greater than 60%, preferably greater than 70% and most preferably greater than 90% by weight relative to the total mass of said effluent.

[0074] Step b) of separation can advantageously be carried out by any method known to a person 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 high pressure and / or low pressure stripping.

[0075] Step b) also allows the separation of at least one aqueous liquid effluent, preferably water. The removal of at least some of the water, and preferably all of it, can be carried out by any methods and techniques known to those skilled in the art. Preferably, the water removal is carried out by settling in a separatory vessel, by drying, by passing through a desiccant, by flash drying, or by a combination of at least two of these techniques. The atomic oxygen content of the hydrocarbon liquid effluent containing the paraffinic hydrocarbons 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, most preferably less than 6,000 ppm, most preferably less than 1,000 ppm by weight, and most preferably less than 500 ppm by weight.The atomic oxygen content in ppm weight in said hydrocarbon liquid effluent is measured by the infrared absorption technique such as, for example, the technique described in patent application US2009 / 0018374A1.

[0076] In a preferred embodiment, at least part of the hydrocarbon liquid effluent consisting of n-paraffins from said step a) is recycled to hydrotreating step a), so that the recycle rate, i.e. the mass ratio between the flow of said recycled liquid effluent and the feed flow introduced into hydrotreating step a), is less than or equal to 2, preferably less than or equal to 1.7, preferably less than or equal to 1.5.

[0077] It is known that using a high liquid recycle rate allows for better exothermic control and, in particular, maintains the temperature difference between the outlet and inlet temperatures of each catalytic zone within a range acceptable for industrial operation of the process. Surprisingly, despite the implementation of a low liquid recycle rate, the present invention allows for optimized exothermic control within the different catalytic zones thanks to the combination of said low recycle rate and a high hydrogen flow rate at the inlet of the first bed.

[0078] Furthermore, the use of a low-consumption recycled liquid facilitates the revamping of existing units. Revamping, in Anglo-Saxon terminology, refers to the revision of the design of equipment already in operation in order to increase its production, technical, economic and environmental performance, as well as its reliability. Step c)

[0079] According to the invention, the process comprises a first step c) of hydroconversion of at least a part and preferably all of the hydrocarbon liquid effluent from step b) of the process in the presence of a first catalyst bifunctional fixed-bed hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one metal from Group VIII and / or at least one metal from Group VIB of the periodic table alone or in mixture, an acidic solid support comprising a silica-alumina or one or more zeolites and optionally at least one binder, said first hydroconversion step being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at a space-hour velocity between 0.1 and 10 h 1 and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm3 / m3 of feed, and preferably between 150 and 750 Nm3 / m3 of feed.

[0080] The operating conditions of the first step c) of hydroconversion are adjusted to favor hydroisomerization or hydrocracking reactions as required. Preferably, the hydroconversion step c) advantageously operates at a temperature between 250 and 450°C, and most preferably between 250 and 400°C, at a pressure between 2 and 10 MPa and most preferably between 3 and 9 MPa, at an hourly volumetric rate advantageously between 0.2 and 7 h⁻¹ and most preferably between 0.5 and 5 h⁻¹, at a hydrogen flow rate such that the hydrogen / feed volume ratio is advantageously between 100 and 1000 Nm³ / m³ of feed and most preferably between 150 and 1000 Nm³ / m³ of feed.

[0081] According to the invention, the catalyst used in step c) is a bifunctional catalyst comprising a hydrogenating phase comprising at least one metal from group VIII and / or at least one metal from group VIB of the periodic table alone or in mixture, a support an acidic solid support comprising a silica-alumina or one or more zeolites and optionally at least one binder.

[0082] The hydro / dehydrogenating function

[0083] The metals of group VIII are advantageously chosen from iron, cobalt, nickel, platinum, and palladium, taken alone or in mixture, and preferably from nickel, cobalt, platinum, and palladium.

[0084] The metals of group VIB are chosen from tungsten and molybdenum, taken alone or in mixture.

[0085] In the case where the Group VIII metals are chosen from among the non-noble metals, the following metal combinations are preferred: nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, cobalt-tungsten, and most preferably: nickel-molybdenum, nickel-tungsten. It is also possible to use combinations of three metals such as, for example, nickel-cobalt-molybdenum or nickel-molybdenum-tungsten. The Group VIII and / or Group VIB metals are introduced by any method known to those skilled in the art, for example, by dry impregnation of the The support is made from metallic precursor(s) solubilized in a solvent, which may be water. One or more organic compounds may also be added during this impregnation step or in a subsequent impregnation step. These organic compounds may contain oxygen and / or nitrogen and / or sulfur. For example, when using oxygenated compounds, examples include compounds comprising one or more functional groups selected from among a carboxyl group, alcohol, ether, ketone, ester, or carbonate, as well as furanic compounds or sugars.

[0086] The content of the catalyst in non-noble group VIII metal is advantageously between 0.5% and 10% by weight of oxide relative to the total weight of said catalyst, preferably between 1% and 8% by weight of oxide and most preferably between 1.5% and 6% by weight of oxide.

[0087] In the case where the metals of group VIII are chosen from among the noble metals, the content of noble metal of group VIII, and preferably the platinum content, in the catalyst used in step c) is between 0.01% and 4% weight, preferably between 0.05% and 2% weight, relative to the total weight of said catalyst.

[0088] The content of the catalyst in metal of group VIB is advantageously between 1% and 50% by weight of oxide relative to the total weight of said catalyst, preferably between 10% and 40% by weight of oxide, most preferably between 15% and 35% by weight of oxide.

[0089] The element content is perfectly measured using X-ray fluorescence.

[0090] The catalyst used in step c) may also advantageously comprise at least one additional metal selected from the group formed by the metals of groups IIIA, IVA and VIIB of the periodic table of elements and preferably selected from gallium, indium, tin and rhenium. Said additional metal is preferably selected from indium, tin and rhenium.

[0091] Preferably, the content of at least one additional metal in the catalyst used in step c) is between 0.01% and 2% by weight, preferably between 0.05% and 1% by weight, relative to the total weight of said catalyst.

[0092] In the case where the first hydroconversion catalyst comprises at least one metal from group VIII and at least one metal from group VIB, said first catalyst is in sulfide form.

[0093] In the case where the first hydroconversion catalyst comprises a noble metal from group VIII, said first catalyst is in reduced form.

[0094] In a preferred embodiment, the catalyst comprises at least one metal from group VIII and preferably nickel and at least one metal from group VIB and preferably tungsten, preferably active in their sulfide form.

[0095] In another embodiment, the catalyst comprises at least one Group VIII metal, and preferably a Group VIII noble metal selected from platinum and palladium. Preferably, the Group VIII metal of the catalyst used in step c) is platinum, preferably active in its reduced form.

[0096] The metallic function is advantageously introduced into the catalyst by any method known to those skilled in the art, such as co-mixing, dry impregnation or exchange impregnation.

[0097] In one embodiment, the sulfur content in the hydroconversion catalyst comprising a Group VIII noble metal is such that the ratio of the number of moles of sulfur to the number of moles of at least one Group VIII noble metal is between 0.3 and 3. In one or more embodiments, the presence of sulfur in the catalyst originates from an optional sulfidation step of the hydroconversion catalyst. In one or more embodiments, the presence of sulfur in the catalyst originates from potentially present impurities, such as, for example, in the alumina binder. In another embodiment, the catalyst does not contain sulfur.

[0098] The acid function.

[0099] According to the invention, the catalyst used in step c) comprises at least one acid support comprising a silica-alumina and / or one or more zeolites.

[0100] In the case where the acidic support comprises and preferably consists of one or more zeolites, said zeolites are selected from structural type zeolites FAU, *BEA, ISV, IWR, IWW, MEI, UWY, taken alone or in mixtures, and preferably selected from structural type zeolites FAU and *BEA, taken alone or in mixtures. In a preferred embodiment, the zeolite is selected from zeolite Y and beta zeolite taken alone or in mixtures, and preferably the zeolite is zeolite Y, and most preferably dealuminated zeolite USY.

[0101] Preferably, the catalyst used in step c) comprises a zeolite content of between 1% and 90% by weight, preferably between 3% and 80% by weight, and more preferably between 4% and 60% by weight, preferably between 4% and 30% by weight and even more preferably between 4% and 20% by weight relative to or total weight of said catalyst.

[0102] Preferably, the acid support of the catalyst used in step c) comprises and is preferably made of silica-alumina.

[0103] The binder

[0104] Optionally, the catalyst support used in step c) may also include a binder. Preferably, the support includes a binder when the support includes a zeolite. The binder is advantageously selected from silica (SiO2), alumina (Al2O3), clays, titanium dioxide (TiO2), and boron dioxide (B2O). 3) and zirconia (ZrO2) alone or in mixtures. Preferably, said binder is chosen from silica and alumina and even more preferably, said binder is alumina in all its forms known to those skilled in the art, such as, for example, gamma alumina.

[0105] A preferred hydroconversion catalyst used in the first hydroconversion step c) comprises and preferably consists of at least tungsten and / or molybdenum and at least nickel and / or cobalt, and preferably nickel and tungsten, and silica-alumina as an acid support, without any other binder. In this case, said catalyst is in sulfide form.

[0106] In this case, said preferred catalyst used in the first step c) of hydroconversion comprises a tungsten and / or molybdenum content advantageously, in oxide equivalent, of between 5% and 50% by weight relative to the finished catalyst, preferably between 10% and 40% by weight and most preferably between 15% and 35% by weight, and a nickel and / or cobalt content advantageously, in oxide equivalent, of between 0.5% and 10% by weight relative to the finished catalyst, preferably between 1% and 8% by weight and most preferably between 1.5% and 6% by weight, and a particular silica-alumina, said silica-alumina having:

[0107] - of alumina and silica with a silica (SiO2) mass content greater than 5% by weight and less than or equal to 95% by weight, preferably between 10 and 80% by weight, preferably a silica content greater than 20% by weight and less than 80% by weight, and even more preferably greater than 25% by weight and less than 75% by weight, the silica content advantageously being between 10 and 50% by weight

[0108] - a specific BET surface area of ​​100 to 500 m2 / g, preferably between 200 m2 / g and 450 m2 / g, and preferably between 200 m2 / g and 300 m2 / g,

[0109] - an average mesopore diameter measured by mercury porosimetry including between 3 and 12 nm, preferably between 3 nm and 11 nm and most preferably between 4 nm and 10.5 nm,

[0110] - a total pore volume measured by mercury porosimetry of between 0.4 and 1.2 ml / g, preferably between 0.4 and 1.0 ml / g and most preferably between 0.4 and 0.8 ml / g,

[0111] - a macropore volume, the diameter of which is greater than 50 nm, less than 0.002 ml / g.

[0112] The average diameter of the mesopores is defined as the diameter corresponding to the cancellation of the curve derived from the mercury intrusion volume obtained from the mercury porosity curve for pore diameters between 2 and 50 m.

[0113] Preferably, the metal distribution coefficient of said preferred catalyst is greater than 0.1, preferably greater than 0.2, and most preferably greater than 0.4. The distribution coefficient represents the distribution of the metal within the catalyst grain. The metal distribution coefficient can be measured by a Castaing microprobe.

[0114] Another preferred hydroconversion catalyst used in said first step c) of hydroconversion comprises and is preferably made of at least one noble metal, said noble metal being platinum, and a silica-alumina as an acid support, without any other binder.

[0115] In this case, said catalyst is in reduced form.

[0116] Preferably, said catalyst comprises between 0.05% and 10% by weight, preferably between 0.1% and 5% by weight of at least one noble metal from Group VIII, preferably selected from platinum and palladium (preferably platinum) deposited on silica-alumina, without any other binder, containing an amount of silica (SiO2) between 1% and 95%, expressed as a percentage by weight, preferably between 5% and 95%, preferably between 10% and 80%, and most preferably between 20% and 70%, and even more preferably between 22% and 45%, said catalyst having:

[0117] - a specific BET surface area of ​​100 to 500 m2 / g, preferably between 200 and 450 m² / g and preferably between 200 and 300 m² / g,

[0118] - an average mesopore diameter measured by mercury porosimetry including between 4 and 12 nm, preferably between 4 and 11 nm and most preferably between 5 and 11 nm,

[0119] - a total pore volume measured by mercury porosimetry of between 0.2 and 1.2 ml / g, preferably between 0.3 and 1.0 ml / g and most preferably between 0.3 and 0.9 ml / g,

[0120] - a macropore volume measured by mercury porosimetry, the diameter of which is greater than 50 nm, less than 0.02 ml / g.

[0121] - an alkali or alkaline-earth compound content of less than 300 ppm by weight and preferably less than 200 ppm by weight.

[0122] The average diameter of the mesopores is defined as the diameter corresponding to the cancellation of the curve derived from the volume of mercury intrusion obtained by mercury porosimetry for pore diameters between 3.7 and 50 nm.

[0123] Preferably, the dispersion of the noble metal of said preferred catalyst is advantageously between 5% and 100%, preferably between 5% and 90%, and most preferably between 10% and 90%. The dispersion, representing the fraction of metal accessible to the reactant relative to the total amount of metal in the catalyst, is advantageously measured, for example, by H2 / O2 titration or by transmission electron microscopy.

[0124] Preferably, the noble metal distribution coefficient of said preferred catalyst is greater than 0.1, preferably greater than 0.2, and most preferably greater than 0.4. The noble metal distribution represents the distribution of the metal within the catalyst grain, the metal being either well or poorly dispersed. Thus, it is possible to obtain platinum that is poorly distributed (for example, detected in a ring whose thickness is significantly less than the radius of the grain) but well dispersed, meaning that all the platinum atoms located in the ring will be accessible to the reactants. The noble metal distribution coefficient can be measured by a Castaing microprobe.

[0125] Most preferably, said preferred hydroconversion catalyst used in said first hydroconversion step c) is a catalyst comprising and preferably made of platinum and a silica alumina support, preferably used in reduced form.

[0126] Preferably, the catalyst used in the first step (c) is shaped into cylindrical or multilobed extrudates such as bilobed, trilobed, or multilobed straight or twisted shapes. According to one or more embodiments, the catalyst used in step (c) is shaped into crushed powders, tablets, rings, beads, or wheels. Other techniques besides extrusion, such as pelletizing or coating, may advantageously be used.

[0127] Preferably, when said catalyst used in step c) contains a noble metal, said noble metal can advantageously be reduced. One preferred method for conducting 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 consist of a two-hour holding period at 150°C followed by a temperature ramp-up to 450°C at a rate of 1°C / min and then a two-hour holding period at 450°C; during the reduction step, the hydrogen flow rate may be 1000 normal m³ hydrogen / m³ catalyst and the total pressure may be maintained 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 unit itself 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).

[0128] Preferably, when said catalyst used in step c) does not contain noble metal, said non-noble metal(s) may advantageously be sulfided by any method known to those skilled in the art using a sulfidating agent, by in-situ sulfidation in the unit or by ex-situ sulfidation before loading into the unit. Step d)

[0129] According to the invention, the process includes a second step d) of hydroconversion of all the hydrocarbon effluent from step c) in the presence of a second fixed-bed hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one metal from group VIII and / or group VIB of the periodic table and a support comprising at least one zeolite selected from the structural type zeolites FER, EUO, *MRE, MTW, MOR, and IZM-2 zeolite alone or in mixture, and at least one binder. The said second hydroconversion stage is carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at a space-hour speed 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 Nm3 / m3 of charge.

[0130] Preferably, no separation step is implemented between hydroconversion steps c) and d). All of the effluent from the first hydroconversion step c) is sent to the second hydroconversion step d).

[0131] Steps c) and d) can advantageously be carried out in a fixed bed within the same reactor, the first and second hydroconversion catalysts then being able to be implemented in one or more catalytic beds, and preferably two separate catalytic beds.

[0132] The operating conditions of the second hydroconversion step d) are adjusted to favor hydroisomerization or hydrocracking reactions as required. Preferably, the hydroconversion step d) advantageously operates at a temperature between 250 and 450°C, and most preferably between 250 and 400°C, at a pressure between 2 and 10 MPa and most preferably between 3 and 9 MPa, at an hourly volumetric rate advantageously between 0.2 and 7 h⁻¹ and most preferably between 0.5 and 5 h⁻¹, at a hydrogen flow rate such that the hydrogen / feed volume ratio is advantageously between 100 and 1000 Nm³ / m³ of feed and most preferably between 150 and 1000 Nm³ / m³ of feed.

[0133] The hydro / dehydrogenating function

[0134] The metals of group VIII are advantageously chosen from iron, cobalt, nickel, platinum, and palladium, taken alone or in mixture, and preferably from nickel, cobalt, platinum, and palladium.

[0135] The metals of group VIB are selected from tungsten and molybdenum, taken alone or in mixture.

[0136] In the case where the metals of Group VIII are chosen from among the non-noble metals, the following metal combinations are preferred: nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, cobalt-tungsten, and most preferably: nickel-molybdenum, nickel-tungsten. It is also possible to use combinations of three metals such as for example nickel-cobalt-molybdenum or nickel-molybdenum-tungsten.

[0137] The metals of Group VIII and / or Group VIB are introduced by any method known to those skilled in the art, for example, by dry impregnation of the substrate with the metal precursor(s) dissolved in a solvent, which may be water. One or more organic compounds may also be added during this impregnation step or in a subsequent impregnation step. These organic compounds may contain oxygen and / or nitrogen and / or sulfur. For example, in the case of using oxygenated compounds, examples include compounds comprising one or more functional groups selected from among a carboxyl group, alcohol, ether, ketone, ester, or carbonate, or even furanic compounds or sugars.

[0138] The content of the catalyst in non-noble group VIII metal is advantageously between 0.5% and 8% by weight of oxide relative to the total weight of said catalyst, preferably between 0.5% and 6% by weight of oxide and most preferably between 1% and 4% by weight of oxide.

[0139] In the case where the metals of group VIII are chosen from among the noble metals, the content of group VIII noble metal, and preferably the platinum content, in the catalyst used in step d) is between 0.01% and 4% weight, preferably between 0.05% and 2% weight, relative to the total weight of said catalyst.

[0140] The content of the catalyst in metal of group VIB is advantageously between 1% and 50% by weight of oxide relative to the total weight of said catalyst, preferably between 10% and 40% by weight of oxide, most preferably between 15% and 35% by weight of oxide.

[0141] The catalyst used in step d) may also advantageously comprise at least one additional metal selected from the group formed by the metals of groups IIIA, IVA and VIIB of the periodic table of elements and preferably selected from gallium, indium, tin and rhenium. Said additional metal is preferably selected from indium, tin and rhenium.

[0142] Preferably, the content of at least one additional metal in the catalyst used in step d) is between 0.01% and 2% by weight, preferably between 0.05% and 1% by weight, relative to the total weight of said catalyst.

[0143] In a preferred embodiment, the catalyst comprises at least one metal from group VIII and preferably nickel and at least one metal from group VIB and preferably tungsten, preferably active in their sulfide form.

[0144] In the case where said catalyst comprises at least one metal from Group VIII and at least one metal from Group VIB, said catalyst may also advantageously contain a doping element selected from phosphorus and boron, taken alone or in combination The mixture, preferably with phosphorus, can be introduced into the matrix or, preferably, deposited on the support. Silicon can also be deposited on the support, alone or with phosphorus and / or boron and / or fluorine.

[0145] The weight content of oxide of said doping element in weight percentage relative to the total weight of said catalyst is advantageously less than 20% and preferably less than 10% and is advantageously at least 0.001%.

[0146] In 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 at least one Group VIII noble metal is between 0.3 and 3. In one or more embodiments, the presence of sulfur in the catalyst originates from an optional sulfidation step of the hydroconversion catalyst. In one or more embodiments, the presence of sulfur in the catalyst originates from potentially present impurities, such as, for example, in the alumina binder. In another embodiment, the catalyst does not contain sulfur.

[0147] In another embodiment, the catalyst comprises at least one metal from Group VIII, and preferably a noble metal from Group VIII selected from platinum and palladium. Preferably, the Group VIII metal of the catalyst used in step c) is platinum, preferably active in its reduced form.

[0148] The metallic function is advantageously introduced into the catalyst by any method known to those skilled in the art, such as co-mixing, dry impregnation or exchange impregnation.

[0149] The acid function.

[0150] According to the invention, the catalyst contains at least one zeolite selected from the following structural type FER zeolites, preferably selected from ZSM-35 and ferrierite, taken alone or in mixture, EUO, preferably selected from EU-1 and ZSM-50, taken alone or in mixture, *MRE, preferably selected from ZSM-48, ZBM-30, EU-2 and EU-11, taken alone or in mixture, MTW, preferably selected from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or in mixture, MOR, preferably selected from mordenite or LZ-211, taken alone or in mixture, and IZM-2 zeolite, taken alone or in mixture. Structural codes are defined in the International Zeolite Association (IZA: http: / / www.iza-structure.org / databases / ) classification.

[0151] Zeolite may also be IZM-2, the structural code of which is not known.

[0152] Preferably, the catalyst contains at least one zeolite selected from IZM-2 zeolite and one MTW structural type zeolite preferably selected from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or in mixture.

[0153] Preferably, the catalyst contains at least one zeolite selected from IZM-2 and ZSM-12, taken alone or in mixture.

[0154] The zeolites are preferably in their acidic form, that is to say, the atomic ratio between the monovalent compensating 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 most 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 acidic form of said zeolites.

[0155] According to a preferred embodiment, the catalyst comprises a support comprising IZM-2 zeolite alone or a support comprising ZSM-12 zeolite alone.

[0156] Zeolite IZM-2 is a crystalline microporous solid whose crystal structure and preparation process are described in patent application FR2918050A1. The structural code of zeolite IZM-2 is not yet known. Zeolite ZSM-12 is a crystalline 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 MRs, and its structural code is MTW. A preparation process for this zeolite is described, for example, in the article "Synthesis of zeolite ZSM-12 in the System (MTEA)2O-Na2O-SiO2-Al2O3-H2O" by S. Ernst et al. (Zeolites, 7, 5, 458-462, DOI10.1016 / 0144-2449(87)90015-7).

[0157] Preferably, the catalyst used in step e) comprises a zeolite content of between 1% and 90% by weight, preferably between 3% and 80% by weight, and more preferably between 4% and 60% by weight, preferably between 4% and 30% by weight and even more preferably between 4% and 20% by weight relative to or total weight of said catalyst.

[0158] The binder.

[0159] Preferably, the binder of the catalyst support in step e) is amorphous or crystalline. Preferably, the binder used in the catalyst support in step e) is advantageously chosen from the group consisting of alumina, silica, silica-alumina, clays, titanium oxide, boron oxide, zirconia, and aluminates, alone or in mixtures. 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, and delta aluminas.

[0160] Preferably, the catalyst used in step e) comprises a binder content of between 10% and 99% by weight, relative to or total weight of said catalyst i.e., so as to ensure the complement to 100% by weight of the elements constituting the catalyst used in step e).

[0161] 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 binder precursor, such as, for example, boehmite, which, upon heat treatment, is transformed into alumina.

[0162] A preferred catalyst for step d) comprises and is preferably made of platinum, and a support comprising and preferably made of a ZSM-12 zeolite and an alumina binder. In this case, said catalyst is in reduced form.

[0163] Another preferred catalyst for step d) comprises and is preferably made of platinum, and a support comprising and preferably made of an IZM-2 zeolite and an alumina binder. In this case, said catalyst is in reduced form.

[0164] In this case, the catalyst used in the second hydroconversion step c) comprises more particularly, and preferably consists of:

[0165] - from 1% to 90% by weight, preferably from 3% to 80% by weight and even more preferred 4% to 60% zeolite by weight, preferably ZSM-12 or IZM-2;

[0166] - from 0.01% to 4% by weight, preferably from 0.05% to 2% by weight of at least one metal of the group VIIIB, preferably platinum;

[0167] - possibly from 0.01% to 2% by weight, preferably from 0.05% to 1% by weight of less one additional metal chosen from the group formed by the metals of groups IIIA, IVA and VIIB;

[0168] - possibly a sulfur content, preferably such that the ratio of the number of moles of sulfur divided by the number of moles of metal(s) in group VIIIB should be between 0.3 and 3; and

[0169] - possibly at least one binder, preferably alumina, ensuring the complement at 100% in the catalyst, relative to the total weight of the catalyst in step d).

[0170] Another preferred catalyst for step d) comprises and is preferably made of nickel and tungsten, and a support comprising and preferably made of ZSM-12 zeolite and an alumina binder. In this case, said catalyst is in sulfide form.

[0171] Another preferred catalyst for step d) comprises and is preferably made of nickel and tungsten, and a support comprising and preferably made of an IZM-2 zeolite and an alumina binder. In this case, said catalyst is in sulfide form.

[0172] In this case, the catalyst used in the second hydroconversion step c) comprises more particularly, and preferably consists of:

[0173] - from 1% to 90% by weight, preferably from 3% to 80% by weight and even more preferred 4% to 60% zeolite weight;

[0174] - of 0.5% and 10% by weight relative to the total weight of said catalyst, in such a way preferred between 1% and 8% by weight and very preferably between 1.5% and 6% by weight of at least one metal from group VIIIB, preferably nickel;

[0175] - 1% and 50% by weight of oxide relative to the total weight of said catalyst, in such a way preferably between 10% and 40% by weight of oxide, most preferably between 15% and 35% by weight of oxide, of at least one metal from group VIB, preferably tungsten;

[0176] - possibly a weight content of oxide of at least one chosen dopant element among phosphorus and boron, taken alone or in mixture, and preferably phosphorus as a percentage by weight relative to the total weight of said catalyst less than 20% and preferably less than 10% and at least 0.001%,

[0177] - possibly from 0.01% to 2% by weight, preferably from 0.05% to 1% by weight of less one additional metal chosen from the group formed by the metals of groups IIIA, IVA and VIIB;

[0178] - possibly at least one binder, preferably alumina, ensuring the complement at 100% in the catalyst, relative to the total weight of the catalyst in step d).

[0179] Preferably, the catalyst used in step d) is shaped into cylindrical or multilobed extrudates such as bilobed, trilobed, or multilobed straight or twisted shapes. According to one or more embodiments, the catalyst used in step e) is shaped into crushed powders, tablets, rings, beads, or wheels. Techniques other than extrusion, such as pelletizing or coating, may advantageously be used. The shaping may also advantageously be carried out in the presence of the various constituents of the catalyst and extrusion of the resulting mineral paste by pelletizing, shaping into beads using a rotary dripper or drum, drop coagulation, oil-drop, oil-up, or any other known process for agglomerating a powder containing alumina and possibly other ingredients selected from those mentioned above.

[0180] Furthermore, the supports used in the process according to the present invention may advantageously have been treated, as is well known to those skilled in the art, with additives to facilitate shaping and / or improve the final mechanical properties of the supports. Examples of additives include, in particular, cellulose, carboxymethyl cellulose, carboxyethyl cellulose, tall oil, xanthan gums, surfactants, flocculating agents such as polyacrylamides, carbon black, starches, stearic acid, polyacrylic alcohol, polyvinyl alcohol, biopolymers, glucose, polyethylene glycols, etc.

[0181] Extrusion can advantageously be carried out using any conventional, commercially available tool. The dough resulting from the mixing process is advantageously extruded through a die, for example using a piston or a single or double screw extrusion die. This extrusion step can advantageously be carried out by any method known to those skilled in the art.

[0182] The catalyst support implemented in step e) according to the present invention is then advantageously subjected to a drying step carried out according to any technique known to a person skilled in the art.

[0183] Preferably, drying is carried out under an airflow. Drying may also advantageously be carried out under a flow of any oxidizing, reducing, or inert gas. Preferably, drying is advantageously carried out between 50 and 180°C, most preferably between 60 and 150°C, and most preferably between 80 and 130°C.

[0184] Said support, possibly dried, then preferably undergoes a calcination step.

[0185] The calcination step is advantageously carried out in the presence of molecular oxygen, for example by performing an air purge, at a temperature advantageously above 200°C and less than or equal to 1100°C. The calcination step can advantageously be carried out in a flow bed, a lick bed, or in a static atmosphere. For example, the furnace used can be a rotary kiln or a vertical radial flow bed kiln. Preferably, the calcination step is carried out for more than one hour at 200°C and less than one hour at 1100°C. The calcination can advantageously be carried out in the presence of steam and / or in the presence of an acidic or basic vapor. For example, the calcination can be carried out under partial pressure of ammonia.

[0186] Post-calcination treatments may optionally be carried out in order to improve the properties of the support, for example textural.

[0187] Preferably, the noble metal contained in said catalyst used in step d) can advantageously be reduced. One preferred method 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 consist of a two-hour hold at 150°C followed by a temperature ramp up to 450°C at a rate of 1°C / min and then a two-hour hold at 450°C; during the reduction step, the hydrogen flow rate may be 1000 normal m³ hydrogen / m³ catalyst and the total pressure may be maintained constant at 0.1 MPa. Any reduction method may advantageously be considered, either in situ (the reduction of the catalyst is carried out in the same unit where the catalytic reaction is performed), 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).

[0188] Preferably, when said catalyst used in step c) does not contain a noble metal, said non-noble metal(s) may advantageously be sulfided by any method known to those skilled in the art using a sulfurizing agent, by in-situ sulfation in the unit or by ex-situ sulfation before loading into the unit.

[0189] In a highly preferred embodiment, the first hydroconversion step c) is carried out in the presence of a first catalyst comprising and preferably consisting of at least one noble metal, said noble metal being platinum, and a silica-alumina as an acid support, without any other binder, said first catalyst being in reduced form, and the second hydroconversion step d) is carried out in the presence of a second catalyst comprising and preferably consisting of platinum, and a support comprising a ZSM-12 zeolite and an alumina binder, in reduced form.

[0190] The proportion of the first hydroconversion catalyst advantageously represents between 10 and 90% and preferably between 20 and 80% of the total catalyst volume, the total catalyst volume being the sum of the volume of the first hydroconversion catalyst and the volume of the second hydroconversion catalyst.

[0191] These proportions are applicable whether the hydroconversion step c) and step d) are carried out in separate reactions or in the same reactor. Step e)

[0192] According to the invention, the process includes a step e) of separating at least a part and preferably all of the effluent from step d).

[0193] Said step e) allows the separation of at least one so-called light gaseous fraction and at least one hydrocarbon liquid effluent. Optionally, step e) also allows the separation of at least a portion of the wastewater and preferably all of the wastewater.

[0194] Said light gaseous fraction comprises at least the hydrogen not converted by the reactions carried out in steps c) and d) and at least a portion of the C4 cracking products, i.e., C1 to C4 compounds 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 possibly an aqueous liquid effluent containing a small amount of residual water (residual water generated either by the hydrodeoxygenation of residual oxygenated compounds present in the feed of steps c) and d) or by slight entrainment of water from step b) to steps c) and d)).

[0195] The separation step e) described can advantageously be carried out by any method known to those skilled in the art, such as, for example, the combination of one or more high- and / or low-pressure separating flasks operated hot or cold, and / or high-pressure and / or low-pressure stripping. This step e) does not require the use of a distillation column. This step e) is advantageously not carried out in a distillation column. Step f)

[0196] According to the invention, the process includes a step f) of stabilizing at least part and preferably all of the hydrocarbon liquid effluent from step e).

[0197] The stabilization step f) can be implemented by any method known to the person skilled in the art, such as, for example, a stripping step advantageously using steam and / or separation within separation balloons and / or a combination of these options.

[0198] Preferably, said stabilization step is a stripping step of at least part and preferably all of the hydrocarbon liquid effluent from step e).

[0199] According to the invention, said stripping step allows the separation of a light gaseous fraction, a naphtha hydrocarbon cut and alternatively a diesel cut or a kerosene cut.

[0200] More specifically, the cut point between the middle distillate and naphtha cuts is adjustable in the range of 80-160°C, preferably around 120°C. The naphtha cut thus contains the compounds that have a boiling point below 120°C and the middle distillate cut contains the compounds that have a boiling point above 120°C. In the presence of the specific catalyst sequence used according to the invention, the operating conditions of steps c) and d) of hydroconversion can advantageously be adjusted to produce a middle distillate cut oriented alternatively either towards the production of bio-kerosene conforming to ASTM D7566, or towards the production of a bio-diesel base.

[0201] The valorization of the bio-naphtha cut is not the object of the present invention, but this cut can advantageously be sent to a steam cracking or reforming unit or valorized as a gasoline base in a mixture with other gasoline bases.

[0202] The process according to the invention advantageously does not include a step of hydrocracking part or all of the effluent from step d).

[0203] Preferably, the process according to the invention does not include a final fractionation or distillation step of the effluent from step d) of hydroconversion, after separation of the naphtha cut.

[0204] According to the invention, the process according to the invention does not include recycling in steps c) and / or d) of hydroconversion, of all or part of the hydrocarbon liquid effluents from step d) of hydroconversion, of the effluents from step e) of separation and / or of the effluents from step f) of stabilization. Description of the figures

[0205] [Fig.1]

[0206] Fig. 1 represents the different stages of the renewable kerosene production process comprising one hydrodeoxygenation stage and two hydroconversion stages and using a specific catalyst for each of the hydroconversion stages.

[0207] The feedstock from renewable sources is sent via pipeline 1 mixed with makeup and / or recycled hydrogen (2) to a hydrotreatment unit (a). The hydrotreated effluent from the hydrotreatment unit (a) is withdrawn via pipeline 3 and sent to the three-phase separation unit (b), which separates at least one hydrogen-rich gaseous effluent (4) and at least one hydrocarbon-rich liquid effluent (5). Unit (b) also removes at least some of the water produced by the hydrodeoxygenation reactions, and preferably all of this water (6).

[0208] The hydrocarbon liquid effluent (5) is sent to a first hydroconversion unit c) in the presence of a supplementary and / or recycled hydrogen stream (7) to produce a second effluent (8) which is sent in full to a second hydroconversion unit d) d) in the presence of a supplementary hydrogen stream (7b) to produce an effluent (9) which is sent to a three-phase separation step (e) allowing the separation of a hydrogen-rich gaseous effluent (10) which may also contain light such as the Ci-C4 cut, water (1) and at least one hydrocarbon liquid effluent (11).

[0209] The hydrocarbon liquid effluent (11) from unit e) is sent to a stripping stabilization unit f) allowing the separation of a light gaseous fraction (13), a naphtha hydrocarbon cut (14), and a middle distillate hydrocarbon cut, alternatively biodiesel or biokerosene (15) having an initial boiling point between 100 and 180°C.

[0210] The examples below illustrate the invention without limiting its scope. EXAMPLES

[0211] Example 1: preparation of a hydrotreating catalyst (Cl).

[0212] The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with molybdenum oxide (MoO3) contents of 22 wt%, in oxide of nickel NiO of 4% by weight and of phosphorus oxide P2O5 of 5% by weight relative to the total weight of the finished catalyst.

[0213] Example 2: Preparation of a first stage hydroconversion catalyst according to the invention (C2).

[0214] The silica-alumina powder is prepared according to the synthesis protocol described in patent EPI 415 712A. The quantities of orthosilicic acid and aluminum hydrate are chosen so as to have a composition of 70% by weight of alumina A12O3 and 30% by weight of silica SiO2 in the final solid.

[0215] This mixture is rapidly homogenized in a commercial colloidal mill in the presence of nitric acid so that the nitric acid content of the suspension exiting the mill is 8% relative to the silica-alumina solid. The suspension is then conventionally dried in a spray dryer from 300°C to 60°C. The resulting powder is shaped in a Z-arm in the presence of 8% nitric acid relative to the anhydrous product. Extrusion is carried out by passing the paste through a die with 1.4 mm diameter orifices. The extrudates thus obtained are oven-dried at 140°C and then calcined under a flow of dry air at 550°C, followed by further calcination at 850°C in the presence of steam.

[0216] The characteristics of the support thus prepared are as follows:

[0217] - an average mesopore diameter measured by mercury porosimetry of 7.7 nm,

[0218] - a total pore volume of 0.49 ml / g,

[0219] - a mesoporous volume of 0.47 ml / g,

[0220] - a macropore volume, the diameter of which is greater than 50 nm less than 0.01 ml / g,

[0221] - a BET surface of 240 m2 / g,

[0222] The silica-alumina extrudates are then subjected to a dry impregnation step with an aqueous solution of platinum tetramine nitrate, left to mature in a water-cooled maturing tank for 24 hours at ambient temperature, and then calcined for two hours under dry air in a flow bed at 450°C (temperature ramp of 5°C / min). The platinum content by weight of the finished catalyst after calcination is 0.57%, its dispersion measured by H2 / O2 titration is 26%, and its distribution coefficient measured by Castaing microprobe is 0.87.

[0223] Example 3: Preparation of a second stage hydroconversion catalyst according to the invention (C3). Synthesis of IZM-2 zeolite.

[0224] Zeolite IZM-2 was synthesized in accordance with the teachings of patent FR 2 918 050 B. A colloidal silica suspension known by the trade name Ludox HS-40, marketed by Aldrich, is incorporated in a solution composed of sodium hydroxide (Prolabo), structuring agent 1,6bis(methylpiperidinium)hexane dibromide, aluminum hydroxide (Aldrich), and deionized water. The molar composition of the mixture is as follows: 1 SiO2; 0.0060 Al2O3; 0.1666 Na2O; 0.1666 1,6bis(methylpiperidinium)hexane; 33.3333 H2O. The mixture is vigorously stirred for half an hour. After homogenization, the mixture is transferred to a PARR-type autoclave. The autoclave is heated for 5 days at 170°C with agitation using a rotary spit (30 rpm). The resulting product is filtered, washed with deionized water to achieve 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 in order to eliminate the structuring agent.The calcination cycle consists of heating to 200°C, holding at this temperature for two hours, heating to 550°C followed by an eight-hour holding period at this temperature, and finally a return to room temperature. The heating is carried out at a rate of 2°C / min. The resulting solid is then refluxed for two hours in an aqueous ammonium nitrate solution (10 mL of solution per gram of solid, ammonium nitrate concentration of 3 M) to exchange the alkali sodium cations for ammonium ions. This refluxing step is performed four times with fresh ammonium nitrate solution. The solid is then filtered, washed with deionized water, and dried overnight in an oven at 100°C.Finally, to obtain the zeolite in its acidic form (protonated H+), a calcination step is carried out at 550°C for ten hours (temperature ramp of 2°C / min) in a flow bed under dry air (2 normal liters per hour per gram of solid). The solid thus obtained was analyzed by X-ray Diffraction and identified as being composed of IZM-2 zeolite. Characterizations using 27A1 isotope NMR, X-ray fluorescence and ICP methods provide the following results for IZM-2:

[0225] - weight percentage of A1VI hexacoordinated aluminium atoms: 5%,

[0226] - ratio of the number of moles of silicon divided by the number of moles of aluminum lattice density, in mole / mole, Si / Al: 72,

[0227] - ratio of the number of moles of sodium divided by the number of moles of aluminium network, in mole / mole, Na / Al: 0.03. Preparation of the IZM-2 / alumina support.

[0228] The IZM-2 / alumina support is obtained by mixing and extruding IZM-2 zeolite with a Pural SB3 type alumina gel. The mixed paste is extruded through a 1.8 mm diameter quadrilobe die. After drying overnight in an oven at 110°C, the extrudates are calcined at 500°C for two hours (temperature ramp of 5°C / min) in a flow bed under dry air (2 normal liters per hour per gram of solid). The weight content of IZM-2 zeolite in the support after calcination is 13% by weight.

[0229] Platinum impregnation on the IZM-2 / alumina support.

[0230] Platinum impregnation is carried out by dry impregnation of the support in a dripping container with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)2. After impregnation, the solid is left to mature for at least five hours in laboratory air and then dried overnight in an oven at 110°C. Finally, a calcination step is carried out under a flow of dry air (2 normal liters per hour per gram of solid) in a tubular furnace under the following conditions:

[0231] - temperature rise from ambient to 450°C at 5°C / min;

[0232] - two-hour plateau at 450°C;

[0233] - descent to ambient temperature.

[0234] The Pt content measured by FX on the calcined C3 catalyst is 0.30% by weight, its dispersion measured by H2 / O2 titration is 50%, its distribution coefficient measured by Castaing microprobe is 0.92.

[0235] Example 4: Preparation of a second stage hydroconversion catalyst according to the invention (C4). Zeolite ZSM-12.

[0236] ZSM-12 zeolite is supplied by Zeolyst. The solid was analyzed by X-ray diffraction and identified as being composed of ZSM-12 zeolite. Characterizations using 27Al isotope NMR, X-ray fluorescence, and ICP methods yielded the following results for ZSM-12:

[0237] - weight percentage of A1VI hexacoordinated aluminium atoms: 0%;

[0238] - ratio of the number of moles of silicon divided by the number of moles of aluminum network, in mole / mole, Si / Al: 43;

[0239] - ratio of the number of moles of sodium divided by the number of moles of aluminium network, in mole / mole, Na / Al: 0.009. Preparation of the ZSM-12 / alumina support.

[0240] The ZSM-12 / alumina support is obtained by mixing and extruding ZSM-12 zeolite with a Pural SB3 type alumina gel. The mixed paste is extruded through a 1.8 mm diameter quadrilobe die. After drying overnight in an oven at 110°C, the extrudates are calcined at 500°C for two hours (temperature ramp of 5°C / min) in a flow bed under dry air (2 normal liters per hour per gram of solid). The weight content of ZSM-12 zeolite in the support after calcination is 6 wt%.

[0241] Platinum impregnation on the ZSM-12 / alumina support.

[0242] Platinum impregnation is carried out by dry impregnation of the support in a dripping container with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)2. After impregnation, the solid is left to mature for at least five hours in laboratory air and then dried overnight in an oven at 110°C. Finally, a calcination step is carried out under a flow of dry air (2 normal liters per hour per gram of solid) in a tubular furnace under the following conditions:

[0243] - temperature rise from ambient to 450°C at 5°C / min;

[0244] - two-hour plateau at 450°C;

[0245] - descent to ambient temperature.

[0246] The Pt content measured by FX on the calcined C4 catalyst is 0.30% by weight, its dispersion measured by H2 / O2 titration is 42%, its distribution coefficient measured by Castaing microprobe is 0.89.

[0247] Example 5: Preparation of a second stage hydroconversion catalyst not according to the invention (C5). Zeolite ZSM-23.

[0248] ZSM-23 zeolite is supplied by Zeolyst. The solid was analyzed by X-ray diffraction and identified as being composed of ZSM-23 zeolite. Characterizations using 27Al isotope NMR, X-ray fluorescence, and ICP methods yielded the following results for EU-2:

[0249] - weight percentage of A1VI hexacoordinate aluminum atoms: 0%,

[0250] - ratio of the number of moles of silicon divided by the number of moles of aluminum lattice density, in mole / mole, Si / Al: 24,

[0251] - ratio of the number of moles of sodium divided by the number of moles of aluminium network, in mole / mole, Na / Al: < 0.005. Preparation of the ZSM-23 / alumina support.

[0252] The EU-2 / alumina support is obtained by mixing and extruding ZSM-23 zeolite with a Pural SB3 type alumina gel. The mixed paste is extruded through a 1.8 mm diameter quadrilobe die. After drying overnight in an oven at 110°C, the extrudates are calcined at 500°C for two hours (temperature ramp of 5°C / min) in a flow bed under dry air (2 normal liters per hour per gram of solid). The weight content of ZSM-23 zeolite in the support after calcination is 9% by weight.

[0253] Platinum impregnation on the ZSM-23 / alumina support.

[0254] Platinum impregnation is carried out by dry impregnation of the support in a dripping container with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)2. After impregnation, the solid is left to mature for at least five hours in laboratory air and then dried overnight in an oven at 110°C. Finally, a step of calcination under a flow of dry air (2 normal liters per hour per gram of solid) in a tubular furnace under the following conditions:

[0255] - temperature rise from ambient to 450°C at 5°C / min,

[0256] - two-hour plateau at 450°C,

[0257] - descent to ambient temperature.

[0258] The Pt content measured by FX on the calcined C5 catalyst is 0.32% by weight, its dispersion measured by H2 / O2 titration is 60%, its distribution coefficient measured by Castaing microprobe is 0.94.

[0259] Example 6: Hydrotreating using the catalyst (Cl) of a feedstock from a renewable source according to a process according to the invention

[0260] The hydrotreating catalyst Cl was evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the industrial process according to the invention. The various steps and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are described below.

[0261] In a temperature-controlled reactor designed to ensure isothermal operation and with a fixed bed containing 190 mL of hydrotreating catalyst Cl, the catalyst having been previously sulfurized, pre-refined used cooking oil with a density of 924 kg / m³ and an oxygen content of 11.3% by weight is hydrotreated. The fatty acid distribution, acid value, unsaponifiable content, and triglyceride polymers of the used cooking oil are detailed in Table 1. Prior to the hydrotreating step, the feedstock is treated with dimethyl disulfide to adjust its sulfur content to 50 ppm by weight.

[0262] [Tables 1] Used Cooking Oil Composition in fatty acids % wt 8:0 0.07 12:0 <0.05 14:0 0.19 15:0 <0.05 16:0 8.38 16:1 0.52 17:0 0.06 17:1 <0.05 18:0 3.72 18:1 trans 0.17 18:1 cis 43.21 18:2 trans 0.26 18:2 cis 40.99 18:3 trans 0.06 18:3 cis 0.65 20:0 0.27 20:1 0.27 22:0 0.64 22:1 <0.05 24:0 0.31 24:1 <0.05 Fatty acid content, g / 100 g 95.152 Acid value mg KOH / g 11.6 Unsaponifiable content, g / 100 g 0.65 Triglyceride polymers, % wt 7.0

[0263] Table 1: Characteristics of the renewable feed used cooking oil used as feed for the hydrotreatment step.

[0264] Before hydrotreating the feedstock, the catalyst is sulfided in-situ in the unit with an isane additive containing 2% by weight of dimethyl disulfide, under a total pressure of 5.1 MPa, at a hydrogen / additized diesel fuel ratio of 700 Nm3 per m3. The volume of isane additive per volume of catalyst per hour is set at 1 h1. The sulfidation is carried out for 12 hours at 350°C, with a temperature ramp of 10°C per hour.

[0265] After sulfidation, the operating conditions of the unit are adjusted in order to perform the hydrotreatment of the feed:

[0266] - VVH (charge volume / catalyst volume / hour): 1 h

[0267] - total working pressure: 5.1 MPa,

[0268] - hydrogen / charge ratio: 700 Nm3 of hydrogen / m3 of charge,

[0269] - temperature: 310°C.

[0270] The hydrogen used is supplied by Air Product and has a purity greater than 99.999% by volume.

[0271] Effluent separation step from the hydrotreatment step according to example 6

[0272] The entire hydrotreated effluent from the hydrotreatment step according to Example 6 is separated using a gas / liquid separator to recover a light fraction consisting mainly of hydrogen, propane, water in the form of vapor, carbon oxides (CO and CO2), and ammonia, and a liquid hydrocarbon effluent consisting mainly of linear hydrocarbons. The water present in the liquid hydrocarbon effluent is removed by sedimentation. The liquid hydrocarbon effluent thus obtained contains an atomic oxygen content of less than 80 ppm w / w, said atomic oxygen content being measured by the infrared adsorption technique described in US patent application 2009 / 0018374, and a sulfur content of 2 ppm w / w and a nitrogen content of less than 1 ppm w / w, said nitrogen and sulfur contents being measured respectively by chemiluminescence and UV fluorescence.The density according to ASTM D4052 and the simulated distillation curve according to ASTM D2887 are also measured on the liquid hydrocarbon effluents. The liquid hydrocarbon effluent is mainly composed of paraffins; its normal paraffin composition is measured by gas chromatography coupled with a flame ionization detector. The characteristics obtained on the liquid hydrocarbon effluent are provided in Table 2.

[0273] [Tables2] Liquid hydrocarbon effluent: Oxygen content, ppm <80; Sulfur content, ppm <2; Nitrogen content, ppm <1; Density at 15°C, kg / m³: 791; Simulated Distillation °C: IBP (0.05% wt) 236.7; 10% wt 302.9; 50% wt 321.9 90% wt 324.4 95% wt 331.0 96% wt 366.1 97% wt 404.7 98% wt 457.3 99% wt 469.6 FBP (99.5% wt) 480.6 Distribution of n-Paraffins by carbon number % wt nC8 0.08 nC9 0.07 nCIO 0.09 nCll 0.04 nC12 0.09 nC13 0.11 nC14 0.19 nC15 1.04 nC16 7.10 nC17 11.36 nC18 74.66 nC19 0.58 nC20 0.61 nC21 0.16 nC22 0.68 nC23 0.01 nC24 0.32

[0274] Table 2: Composition of the liquid hydrocarbon effluent used as feedstock for hydroconversion.

[0275] Example 7 not in accordance with the invention: hydroconversion of the liquid hydrocarbon effluent from Example 6 and using only the C4 catalyst for the production of kerosene

[0276] The C4 hydroconversion catalyst was evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the non-conforming industrial process according to the invention. The various steps and unit operations of the pilot unit mimicking the industrial implementation of the non-conforming process according to the invention are described below.

[0277] In a temperature-controlled reactor designed to ensure isothermal operation and with a fixed bed loaded with 50 ml of C4 hydroconversion catalyst, the catalyst being previously activated by reduction, the hydroconversion of the liquid hydrocarbon effluent from Example 6 is carried out.

[0278] The C4 catalyst undergoes an in-situ hydrogen reduction step in the unit, under a total pressure of 5.1 MPa, with a hydrogen / catalyst volume ratio of 700 Nm3 per m3. The reduction is carried out for 2 hours at 400°C, with a temperature ramp of 10°C per hour.

[0279] After reduction, the operating conditions of the unit are adjusted to perform the hydroconversion of the liquid hydrocarbon effluent within the following range of operating conditions:

[0280] - VVH (charge volume / catalyst volume / hour) = 1 h \

[0281] - total working pressure: 5.1 MPa,

[0282] - hydrogen / charge ratio: 350 Nm3 of hydrogen / m3 of charge.

[0283] The hydrogen used and entering the hydroconversion step is supplied by Air Product, it has a purity greater than 99.999% by volume, it is free of hydrogen sulfide.

[0284] At the outlet of the hydroconversion reactor, the reaction effluent is sent to a gas-liquid separation stage using a flash balloon operated at a pressure comparable to that of the hydroconversion reactor. The hydrogen-rich gas phase is sent to the unit's gas outlet. The liquid hydrocarbon phase is depressurized and sent to a stripper to stabilize the hydroconverted liquid effluent; the gas phase collected at the top of the stripper is also sent to the unit's gas outlet.

[0285] At the unit gas outlet, an online gas chromatography analysis and a gas meter allow calculation of the mass of light hydrocarbons produced (primarily hydrocarbons with 1 to 5 carbon atoms) and present in the hydrogen stream. The resulting liquid effluent is weighed separately, topped off at 120°C to remove the naphtha fraction, and the 120°C+ liquid effluent is then reweighed and analyzed, notably by measuring the crystal disappearance point (process operation not conforming to the kerosene target).

[0286] Temperature steps in the range of 250 to 400°C were performed to adjust the severity of the single hydroconversion step. The (typically daily) measurement of the crystal disappearance point (non-compliant process operation in kerosene target) of the 120°C+ liquid effluent allows monitoring of the catalyst performance at each temperature step. For each temperature, the test duration is extended until a stable disappearance point of crystals is obtained (non-compliant operation of the process in kerosene target).

[0287] Once the crystal disappearance point is stable, the 120°C+ (kerosene) cutting yield at the terminals of the non-conforming process is determined according to the following calculation:

[0288] Yield 120°C+ (kerosene) = [(mass of liquid effluent 120°C+) / (mass of feedstock)] x 100, the feedstock here corresponding to the hydrocarbon effluent from example 5.

[0289] The 120°C+ liquid effluent mass corresponds to the quantity of 120°C+ 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 stage during the same period of time.

[0290] The temperature adjustment was carried out so as to achieve a kerosene cut meeting the specifications of ASTM D7566. The characteristics obtained on the kerosene cut according to non-conforming example 7, as well as the operating conditions and the associated kerosene yield are reported in summary table 3.

[0291] Example 8 according to the invention: hydroconversion of the liquid hydrocarbon effluent from Example 6 according to a conforming process, and using a sequence of C2 catalysts followed by a C4 catalyst in the same reactor for the production of kerosene

[0292] In a temperature-controlled reactor designed to ensure isothermal operation and with a fixed bed loaded with 25 ml of hydroconversion catalyst C2 (head of bed) and 25 ml of hydroconversion catalyst C4 (bottom of bed), the catalysts having been previously activated by reduction, the hydroconversion of the liquid hydrocarbon effluent from Example 6 is carried out. The said liquid hydrocarbon effluent is therefore first in contact with the catalyst C2.

[0293] The C2 and C4 catalysts undergo an in-situ hydrogen reduction step in the unit, under a total pressure of 5.1 MPa, with a hydrogen / total catalyst volume ratio of 700 Nm3 per m3. The reduction is carried out for 2 hours at 400°C, with a temperature ramp of 10°C per hour.

[0294] The C2 and C4 hydroconversion catalysts were evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the industrial process according to the invention. The various steps and unit operations of the pilot unit, mimicking the industrial implementation of the process according to the invention, are identical in every respect to those described in Example 7, which is not in accordance with the invention. The characteristics obtained on the kerosene cut according to the compliant Example 8, as well as the conditions operating parameters and associated kerosene yield are reported in summary table 3.

[0295] The conforming process according to Example 8 produces a kerosene cut that meets the specifications of ASTM D7566 with an improved yield compared to that obtained according to the non-conforming Example 7. The yield increase is 13.5% by weight.

[0296] Example 9 according to the invention: hydroconversion of the liquid hydrocarbon effluent from Example 5 according to a non-conforming process, and using a sequence of C2 catalysts followed by a C3 catalyst in the same reactor for the production of kerosene

[0297] The hydroconversion catalysts C2 and C3 were evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the industrial process according to the invention. The various steps and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are identical in all respects to those described in Example 8 according to the invention.

[0298] The conforming process according to Example 9 produces a kerosene cut that meets the specifications of ASTM D7566 with a higher yield than that obtained according to the non-conforming Example 7. The yield increase is 12.9% by weight.

[0299] The use of the process according to the invention, carried out by means of a chain of two specific catalysts C2 then C3 in the same reactor, therefore makes it possible to maintain the overall selectivity of the process in the 120°C+ kerosene cut of interest for said process.

[0300] Example 10 not in accordance with the invention: hydroconversion of the liquid hydrocarbon effluent from Example 5 according to a conforming process, and using a sequence of C2 catalysts followed by a C5 catalyst in the same reactor for the production of kerosene

[0301] The hydroconversion catalysts C2 and C5 were evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the industrial process according to the invention. The various steps and unit operations of the pilot unit, mimicking the industrial implementation of the process according to the invention, are identical in every respect to those described in Example 8, which conforms to the invention. The characteristics obtained on the kerosene cut according to the non-conforming Example 10, as well as the operating conditions and the associated kerosene yield, are reported in Summary Table 3.

[0302] The non-conforming process according to Example 10 using a non-conforming C5 catalyst produces a kerosene cut meeting the specifications of ASTM D7566 with a yield close to that of the non-conforming Example 7, but less favorable than those obtained with the conforming process using C3 or C4 catalysts. The yield loss is 12.9% by weight compared to the conforming process using the C3 catalyst, and 13.5% compared to the conforming process using the C4 catalyst.

[0303] [Tables3] Example 7 Example 8 Example 9 Example 10 ASTM non-compliant compliant non D756S-24a compliant compliant Table A2.1 Catalyst 1st bed C2 C2 C2: WH 1st bed ï 2 A 4 Catalyst 2nd bed: - €4 C3 C5 WH 2nd Ht - 2 1.33 1.33 WH overall ■ #7 1 1 1 1 Total pressure iMPa} 5„1 5.1 5.1 5.1 H2 / HL ratio 35-3 350 350 350 Temperature °C 323 323 330 330 .................................................... Efficiency C1-C4 jgas) k pos 6.0 3.5 3.7 9.9 Efficiency C5-120°C (naphtha) R ntt î 27.1 16.6 16.8 24.5 Tenderness Î2O"C+ (kerosene) f 4pdl) 66.8 80.3 75.7 66.4 - Mass of vaporization at 15°C (g / cm3) 756 763 762 767 730-772 - Crystal disappearance point (°C) <-60 -45 -48 -50 Max -40 - Smoke point (mm) ' 25 >25 >25 >25 Min 25 - Distillation ASTM 086 10% vof ro ​​147 161 162 143 Min 205 T90-T10 (°C) 132 123 125 138 Min : 22 FBP 299 298 297 298 Max 300 Residual (%weight) 1.0 1.0 1.5 1.2 Max 1.5 Losses (%weight) 0.4 0.7 0.4 0.6 Max 1.5

[0304] Summary Table 3

[0305] Example 11 not in accordance with the invention: hydroconversion of the liquid hydrocarbon effluent from example 5 and using only the C4 catalyst for the production of diesel (Pt / ZSM-12)

[0306] The C4 hydroconversion catalyst was evaluated in a pilot unit, the various steps and unit operations of which, mimicking the industrial implementation of the process according to the invention, are identical in every respect to those described in Example 7, which is not in accordance with the invention. The only difference lies in the adjustment of the operating conditions to achieve a limit of filterability (LFR) of -15°C on the 120°C+ cut produced, representative of a "winter diesel" fuel target according to standard NF EN 590. The characteristics obtained on the diesel cut according to the non-conforming Example 11, as well as the Operating conditions and associated diesel cutting yield are reported in summary table 4.

[0307] Example 12 according to the invention: hydroconversion of the liquid hydrocarbon effluent from Example 5 according to a conforming process, and using a sequence of C2 catalysts followed by a C4 catalyst in the same reactor for the production of diesel fuel

[0308] The hydroconversion catalysts C2 and C4 were evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the industrial process according to the invention. The various steps and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are identical in all respects to those described in Example 8, according to the invention.

[0309] The difference lies in the conversion rate of the hydroconversion step, which is adjusted to achieve a limit of filterability (LFR) of -15°C on the 120°C+ cut produced, representative of a "winter diesel" fuel target. The characteristics obtained on the diesel cut according to compliant Example 12, as well as the operating conditions and the associated diesel cut yield, are reported in Summary Table 4.

[0310] The compliant process according to Example 12, using a sequence of catalysts C2 and C4, makes it possible to produce a diesel base oil with desired cold-weather properties, in particular those of the European standard EN590, with an overall process yield equal to that of the non-compliant Example 11. The performance of the compliant process is therefore not degraded compared to that of the non-compliant process for the production of a diesel base oil.

[0311] [Tables4] Example 11 Example 12 Non-compliant Compliant Catalyst 1st bed C4 C2 WH 1st bed (h4) 1 2 Catalyst 2nd bed - C4 WH 2nd bed OH - 2 WH overall 1 1 Total pressure (MPa) 5.1 5.1 H2 / HC ratio (NmVm3) 350 350 Temperature 'C 298 301 C1-C4 efficiency (gas) (%wt) 1.0 0.8 C5-120°C efficiency (naphtha) (%wt) 3.5 3.1 120°C+ efficiency (diesel) (%wt) 95.6 96.2 - Density at 15°C 0.7830 0.7841 filter plugging point temperature (fC) -19 -18 - Distillation ASTM D86 10% vol rcj 268 271 50% vol CQ 290 292 90% vol rn 297 298 FBP ra 321 315 Residue (% wt) 1.0 1.0 Losses (% wt) 0.5 0.7

[0312] Table 4 summary for diesel production.

Claims

1. Demands A process for treating a feedstock from a renewable source to produce, alternatively, a diesel base or a kerosene cut, comprising at least the following steps and preferably consisting of: a) a hydrotreating step of said feed in the presence of a fixed-bed catalyst, said catalyst comprising a hydrogenating function and an oxide support, at a temperature between 200 and 450°C, at a pressure between 1 and 10 MPa, at a spatial velocity between 0.1 and 10 h 1 and in the presence of a total quantity of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 2000 Nm3 of hydrogen / m3 of feed, b) a step of separating at least a portion of the effluent from step a) into at least a light gaseous fraction, at least a hydrocarbon liquid effluent consisting of n-paraffins, and at least an aqueous liquid effluent, (c) a first hydroconversion step of at least a portion of the hydrocarbon liquid effluent from step (b) in the presence of a first bifunctional fixed-bed hydroconversion catalyst, said catalyst comprising at least one metal from Group VIII and / or Group VIB of the periodic table, taken alone or in mixture, and an acidic solid support comprising at least one silica alumina and / or one or more zeolites, said first hydroconversion step being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at a spatial velocity between 0.1 and 10 h⁻¹, and in the presence of a total quantity of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm³ / m³ of feed, d) a second hydroconversion step of all the hydrocarbon effluent from step c) in the presence of a second fixed-bed hydroconversion catalyst, said catalyst comprising a hydrogenating phase comprising at least one metal from Group VIII and / or Group VIB of the periodic table and a support comprising at least one zeolite selected from the structural type zeolites FER, EUO, *MRE, MTW, MOR and ZEOLITH IZM-2 alone or in mixture, and at least one binder, said second step hydroconversion being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at a spatial velocity between 0.1 and 10 h⁻¹, and in the presence of a total quantity of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm³ / m³ of feed; e) a separation step of at least a portion of the effluent from step c) which allows the separation of at least one gaseous fraction and at least one hydrocarbon liquid effluent; f) a stabilization step of at least a portion of the hydrocarbon liquid effluent from step d), so as to separate three hydrocarbon fractions: a light gaseous fraction, a naphtha hydrocarbon fraction, and alternatively a diesel base or a kerosene fraction; no recycling step in hydroconversion steps c) and / or d) of all or part of the hydrocarbon liquid effluents from step d) of hydroconversion,effluents from stage e) separation and / or effluents from stage f) stabilization not being implemented.

2. A process according to claim 1 wherein the feedstock from renewable sources is selected from vegetable oils, animal fats, used cooking oils, oils of microbial origin, fish oils, pine oils, long paraffins (waxes) from the Fischer-Tropsch process, crude or having undergone pretreatment, or mixtures of such feedstocks, containing triglycerides and / or free fatty acids and / or esters.

3. A method according to any one of claims 1 or 2 wherein in step a), the feed is brought into contact with a fixed bed catalyst at a temperature between 220 and 350°C, at a pressure between 1 and 6 MPa, in the presence of hydrogen and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 150 and 1000 Nm3 of hydrogen / m3 of feed.

4. A process according to any one of claims 1 to 3 wherein the hydrotreating catalyst used in step a) comprises a hydro-dehydrogenating function comprising at least one metal of the group VIII and / or of group VIB, taken alone or in 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. A process according to any one of claims 1 to 4 wherein the hydrocarbon liquid effluent consisting of n-paraffins from said step b) comprises a content of compounds boiling at a temperature above 370°C strictly greater than 50%, preferably greater than 60%, preferably greater than 70% and most preferably greater than 90% by weight relative to the total mass of said effluent.

6. A process according to any one of claims 1 to 5 wherein at least a portion of the hydrocarbon liquid effluent consisting of n-paraffins from said step a) is recycled to hydrotreating step a), such that the recycle rate, i.e. the mass ratio between the flow of said recycled liquid effluent and the feed flow introduced into hydrotreating step a), is less than or equal to 2, preferably less than or equal to 1.7, preferably less than or equal to 1.

5.

7. A process according to any one of claims 1 to 6 wherein a hydroconversion catalyst used in the first step c) of hydroconversion comprises nickel and tungsten, and silica-alumina as a support, without any other binder.

8. A process according to any one of claims 1 to 6 wherein a hydroconversion catalyst used in the first step c) of hydroconversion comprises and is preferably made up of at least one noble metal, said noble metal being platinum, and a silica-alumina as an acid support, without any other binder.

9. A process according to any one of claims 1 to 8 wherein the catalyst used in the second hydroconversion step d) contains at least one zeolite selected from IZM-2 zeolite and an MTW structural type zeolite preferably selected from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or in mixture and preferably at least one zeolite selected from IZM-2 and ZSM-12, taken alone or in mixture.

10. A process according to any one of claims 1 to 9, wherein the catalyst used in the second hydroconversion step (d) comprises and is preferably made of platinum, and a support comprising and of preference consisting of a ZSM-12 zeolite and an alumina binder. In this case, said catalyst is in reduced form.

11. A process according to any one of claims 1 to 9 wherein a catalyst used in the second hydroconversion step d) comprises and is preferably made of platinum, and a support comprising and preferably made of an IZM-2 zeolite and an alumina binder, in this case, said catalyst is in reduced form.

12. A process according to any one of claims 1 to 11 wherein the first hydroconversion step c) is carried out in the presence of a first catalyst comprising and preferably consisting of at least one noble metal, said noble metal being platinum, and a silica-alumina as an acid support, without any other binder, said first catalyst being in reduced form and the second hydroconversion step d) is carried out in the presence of a second catalyst comprising and preferably consisting of platinum, and a support comprising a ZSM-12 zeolite and an alumina binder, in reduced form.

13. A method according to any one of claims 1 to 12 wherein the proportion of the first hydroconversion catalyst represents between 10 and 90% and preferably between 20 and 80% of the total catalyst volume, the total catalyst volume being the sum of the volume of the first hydroconversion catalyst and the volume of the second hydroconversion catalyst.

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