Method for producing renewable kerosene comprising two hydroconversion steps and using a specific catalyst in the second hydroconversion step

Abstract

The present invention describes a method for treating a feedstock originating from a renewable source, comprising a hydrotreatment step a), a step b) of separation of the effluent resulting from step a) into a light fraction and a hydrocarbon effluent, a step c) of hydroconversion of the hydrocarbon effluent resulting from step c), using a catalyst comprising a hydrogenating phase that contains at least one metal from group VIII and / or at least one metal from group VIB, and an acid support comprising a silica-alumina or one or more zeolites, a step d) of separation of the effluent resulting from step c), which makes it possible to separate at least one gaseous fraction and at least one hydrocarbon effluent, a step e) of fractionation of the hydrocarbon effluent resulting from step d) into at least one heavy cut having an initial boiling point of between 250°C and 300°C, or a step of stabilisation of the hydrocarbon effluent resulting from step d) in order to separate a heavy cut having an initial boiling point of between 100°C and 180°C, a second hydroconversion step f), distinct from step c), of hydroconversion of the effluent resulting from step e), carried out in the presence of a specific catalyst, a separation step g), and a fractionation or stabilisation step h), wherein steps d) and g), and steps e) and h), are able to be carried out in the same unit or in separate units.

Description

[0001] Description

[0002] Title: PROCESS FOR THE PRODUCTION OF RENEWABLE KEROSENE COMPRISING TWO HYDROCONVERSION STAGES AND USING A SPECIFIC CATALYST IN THE SECOND HYDROCONVERSION STAGE

[0003] Field of invention

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

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

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

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

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

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

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

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

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

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

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

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

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

[0017] The choice of the nature of the hydrogenating function, of the noble metal or sulfide type, depends on different criteria, of an economic nature (the price of noble metals is significantly higher than that of transition metals of groups VI B and VIII) or of a chemical nature (impact of the presence of contaminants). Thus, the hydrogenating activity of noble metals is higher than that of transition metal sulfides when the partial pressure of hydrogen sulfide (H2S) in the reaction medium is low or even zero. Conversely, the hydrogenating activity of transition metal sulfides is higher than that of noble metals when the partial pressure of H2S in the reaction medium becomes significant (C. Marcilly, Acid-Base Catalysis, volume 2, 2003, Technip editions). Patent application US2022 / 0127537 teaches a process for hydrotreating a renewable feedstock.Said method comprises a step of hydrotreating the feedstock in the presence of hydrogen and a hydrotreating catalyst to deoxygenate said feedstock and thus produce a hydrotreated effluent. Said method comprises a step of hydroisomerization, in the presence of hydrogen and a hydroisomerization catalyst, of an effluent from the hydrotreated effluent to obtain a hydroisomerized effluent.In the case of a two-step process, the hydroisomerization catalyst used may comprise a group VIII metal selected from Pt and Pd, alone or in combination and a support which may be amorphous or crystalline selected from alumina, amorphous silica alumina, ferrierite, ALPO-31, SAPO-11, SAPO-31, SAPO, 37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-11, MeAPO-31, MeAPO-41, MgAPSO-11, MgAPSO-31, MgAPSO-41, MgAPSO-46, ELAPSO-11, ELAPSO-3, EMAPSO-41, ELAPSO-11, ELAPSO-31, ELAPSO-41, laumontite, cancrinite, Offretite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.

[0018] In a first embodiment claimed in claim 1, said method comprises a step of separating the hydrotreated effluent and / or the isomerized effluent to obtain a vapor effluent and a liquid effluent; a step of distilling the liquid effluent or the hydroisomerized effluent to produce a kerosene cut and a diesel cut and a step of hydrocracking the diesel cut to obtain a hydrocracked effluent comprising a kerosene cut. The hydrocracking catalyst comprises a metal from group VIII or a metal from group VI B on an amorphous support of silica alumina type or on a zeolitic support. In the case where the hydrocracking catalyst is zeolitic, the active phase comprises a metal from group VIII optionally in combination with a metal from group VI B.The 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. Said process therefore allows the co-production of diesel and kerosene cuts which are separated by a distillation step and the diesel cut obtained then undergoes a hydrocracking step.

[0019] Figure 1 of patent application US2022 / 0127537 describes a process for hydrotreating vegetable oil followed by separation of the hydrotreated effluent in a hot separator to recover a liquid effluent which is hydroisomerized. The hydroisomerized effluent is then fractionated in a distillation column to produce a kerosene cut, a light diesel cut and a heavy diesel cut comprising n-paraffins. Said heavy diesel cut is then sent to a hydrocracking step to convert the heavy diesel cut into kerosene.The hydrocracked effluent can be sent to the hydroisomerization stage at different levels: either directly upstream of the hydroisomerization reactor, i.e. without prior gas / liquid separation, or after prior separation of the effluent by a flash drum producing a gas fraction and a liquid fraction, the liquid fraction of hydrocracked effluent being directed either upstream of the hydroisomerization reactor or downstream. According to an optional embodiment, said process according to Figure 1 therefore comprises a stage of hydrocracking the heavy diesel cut and then recycling it to the hydroisomerization stage. In the present invention, there is no recycling of the effluent from the second hydroconversion stage upstream of the first hydroconversion stage.

[0020] Figure 2 of patent application US2022 / 0127537 describes a process for treating vegetable oil in which the hydrocracking and hydroisomerization steps are carried out successively in the same reactor. The process comprises a vegetable oil hydrotreatment step followed by separation of the hydrotreated effluent in a hot separator to recover a first hydrotreated liquid effluent which is then hydroisomerized in a reactor comprising a hydroisomerization catalyst bed, the latter being positioned downstream of a hydrocracking catalyst bed. The hydrocracking catalyst, positioned upstream of the hydroisomerization catalyst, is charged with the heavy fraction of diesel recycled from the hydroisomerized effluent fractionation step. The hydrotreated liquid effluent is thus sent between the hydrocracking catalyst bed and the hydroisomerization catalyst bed.The hydroisomerized effluent is then fractionated in a distillation column to produce a kerosene cut, a light diesel cut and a heavy diesel cut comprising n-paraffins. Said heavy diesel cut is then sent to the catalytic bed containing the hydrocracking catalyst located upstream of the hydroisomerization catalyst to convert the heavy diesel cut into kerosene. The hydrocracked effluent can be sent to the hydroisomerization stage. In contrast to the present invention, said process according to Figure 1 comprises a stage of hydrocracking the heavy diesel cut before its recycling in the hydroisomerization stage and the hydrocracking and hydroisomerization reactions are carried out successively in the same reactor.

[0021] In a second embodiment claimed in claim 16, said process comprises obtaining a diesel cut from the hydroisomerized effluent, said diesel cut is characterized in that the normal paraffin concentration of a given carbon number is at least twice as high as the corresponding normal paraffin concentration in said hydroisomerized effluent. Said diesel cut undergoes a hydrocracking or hydroisomerization step to obtain a hydrocracked effluent comprising a kerosene cut. Unlike the present invention, said process therefore comprises either a step of hydrocracking the diesel cut obtained, or a step of recycling said diesel cut in the hydroisomerization step.

[0022] In a third embodiment claimed in claim 19, said process comprises a step of separating said hydroisomerized effluent into a vapor effluent and a liquid effluent; a step of distilling the liquid effluent or the hydroisomerized effluent to obtain a kerosene cut and a diesel cut, and a step of hydrocracking said diesel cut to obtain a hydrocracked effluent comprising a kerosene cut. Said process therefore allows the co-production of diesel and kerosene cuts which are separated by a distillation step and the diesel cut obtained then undergoes a hydrocracking step.

[0023] Figure 3 of patent application US2022 / 0127537 describes a process for treating vegetable oil in which the hydrotreatment and hydroisomerization steps are carried out in the same reactor, the catalytic bed comprising the hydrotreatment catalyst being located upstream of the catalytic bed containing the hydroisomerization catalyst. The process comprises a step of hydrotreatment of vegetable oil followed by a step of hydroisomerization of the hydrotreated effluent in a reactor comprising a bed of hydroisomerization catalyst downstream of a bed of hydrotreatment catalyst. The hydroisomerized effluent is then sent to a hydrocracking step, mixed with a heavy diesel fraction produced by fractionation of the hydrocracked effluent in a distillation column. The fractionation makes it possible to produce a kerosene cut, a light diesel cut and a heavy diesel cut comprising n-paraffins.Said heavy diesel cut is then recycled in the hydrocracking stage. Unlike the present invention, said process according to Figure 3 comprises a recycle of the heavy diesel cut as feedstock for the hydrocracking stage in a mixture with the effluent from the hydroisomerization stage carried out in the reactor as hydrotreatment. Unlike the present invention, the hydrotreatment and hydroisomerization stages are carried out successively in the same reactor.

[0024] Figure 4 of patent application US2022 / 0127537 describes a process for treating vegetable oil in which the hydroisomerization and hydrocracking steps are carried out in the same reactor as in Figure 2. The process comprises a vegetable oil hydrotreatment step followed by separation of the hydrotreated effluent in a hot separator and then in a stripping column to recover a light diesel cut (C17-) and a heavy diesel cut (C18+). The heavy diesel cut is sent to a hydrocracking step in a reactor comprising a hydrocracking catalyst bed upstream of a hydroisomerization catalyst bed and the light diesel cut is sent to the hydroisomerization catalyst bed present in the same reactor.The effluent comprising kerosene from said reactor is sent to a stripping column to produce a kerosene cut which is then recycled, only in part, into the catalytic bed comprising the hydroisomerization catalyst. Unlike the present invention, said process according to Figure 4 comprises a step of separating the hydrotreated effluent into light and heavy diesel cuts before sending said light and heavy diesel cuts respectively to the hydroisomerization and hydrocracking catalyst.

[0025] US Patent 8,324,439 B2 teaches a process for treating renewable feedstocks of plant or animal origin. Said process comprises a step of hydrotreating the renewable feedstock, a step of separating the hydrotreated effluent to obtain hydrogen, other gases and at least one effluent containing hydrocarbons. Said process then comprises a step of hydroisomerizing at least a portion of said effluent containing hydrocarbons in the presence of a selective hydroisomerization catalyst, said catalyst comprising at least one one-dimensional zeolite with 10 MR and at least one metal from group VIII and / or group VIB. Finally, said process comprises a step of separating the hydroisomerized effluent to obtain hydrogen, other gases and at least one diesel-type cut.Said catalyst comprises at least one one-dimensional zeolite with 10 MR of structural code TON or EUO, or a zeolite ZSM-48, ZBM-30, IZM-1, COK-7, EU-2 and EU-11, alone or in a mixture. Unlike the process according to the invention, the heavy cut of the hydroisomerized effluent is not sent to a second hydroisomerization step separate from the first.

[0026] Patent application US 2014 / 0291200 A1 describes a process for producing diesel with a sulfur content of less than 10 ppm from renewable feedstocks, comprising the following steps: a) a step of hydrotreating renewable feedstocks, b) a step of separating the hydrotreated effluent from a) into a hydrogen-rich gaseous fraction and a liquid hydrocarbon fraction, c) a step of removing the hLS dissolved in the liquid hydrocarbon effluent from b), d) a step of hydroisomerization of at least a portion of the liquid effluent from step b) c) rather no? on a catalyst comprising at least one metal from group VIII or a combination of a metal from group VIII and a metal from Group VIB on an amorphous support of alumina or silica-alumina type or crystalline, i.e. comprising a 10 MR zeolite chosen from zeolites of structural type TON, EUO or chosen from zeolites ZSM-48, ZBM-30, IZM-1, COK-7, EU-2 and EU-11 alone or as a mixture,e) a step of separating the effluent from d) into a hydrogen-rich gaseous fraction and at least one diesel fraction.,

[0027] The recycling of gaseous fractions from separation steps b) and / or e) in steps a) and / or d) is mentioned with the aim of better control of the exothermicity of the reactions within step a). Unlike the process according to the invention, the heavy cut of the hydroisomerized effluent is not sent to a second hydroisomerization step separate from the first.

[0028] Patent application US 2021 / 0395620 A1 describes a process for producing kerosene from renewable feedstocks. The process comprises a first step of hydrotreating the renewable feedstock in dilution with a hydrocarbon cut, followed by a hydrocracking step. The hydrocracked effluent is then separated into at least one gaseous fraction and at least one liquid fraction. More specifically, the separation step allows the production of a kerosene-type cut and a heavy diesel-type cut. At least a portion of the hydrocracked effluent, more specifically the kerosene-type cut, is then directed to a separate hydrodearomatization step comprising a hydroisomerization catalyst bed and a hydrodearomatization catalyst bed in order to make it compatible with current product specifications, in particular the reduction of the aromatics content for the kerosene-type cut.Unlike the process according to the invention, the kerosene cut is directed towards a second hydroconversion stage (and not the diesel cut), the diesel cut being recycled within the first hydrotreatment stage followed by hydrocracking.

[0029] Patent applications WO 2010 / 000934 A1 and US 2010 / 000908 describe a process for producing kerosene from renewable feedstocks, comprising a first step of hydrodeoxygenation of renewable feedstocks allowing the production of n-paraffins. A second step according to the process allows the hydroisomerization of the n-paraffins from the first step and the production of an effluent rich in iso-paraffins. The latter is sent to a separation and fractionation step allowing the production of a cut boiling in the gasoline range, a cut boiling in the kerosene range and possibly a cut boiling in the diesel range. The cut having a boiling point greater than or equal to 200°C, possibly the cut boiling in the diesel range, is then specifically recycled to a hydroisomerization stage which may be identical to or distinct from the first hydroisomerization stage in order to be reisomerized there.In the case where the second hydroisomerization step is distinct from the first, the effluent from the second hydroisomerization step is returned to the separation and fractionation step as described in Figure 3. The isomerization catalyst(s) comprise a group VIII metal and a support comprising a zeolite chosen from SAPO-11, SAPO-41, ZSM-22, ZSM-23 or ferrierite. In the case where the second hydroisomerization step is distinct from the first, the hydroisomerization catalysts may be identical or different in the two hydroisomerization steps. Preferably, in the first hydroisomerization, the catalyst may be selected so as to be less cracking and in the second hydroisomerization, the second catalyst may be chosen so as to promote the cracking reactions.

[0030] Patent application US 2014 / 0005450 A1 describes a method for manufacturing synthetic distillates, based on the hydrotreatment of a feedstock and then its hydroisomerization / hydrocracking followed by a step of separating the paraffins (n-paraffins and i-paraffins) generated on a molecular sieve. More specifically, the passage on a molecular sieve makes it possible to separate a cut rich in n-paraffins and a cut rich in i-paraffins. The cut rich in i-paraffins is then fractionated by distillation allowing the production of at least one fraction of heavy distillates and at least one fraction of light distillates. A recycle of the n-paraffins separated from the molecular sieve to the hydroisomerization / hydrocracking step is described. This is a selective recycle of a stream according to its chemical family and not according to its distillation interval. A recycle of a portion of the heavy distillate fraction is also described.The hydroisomerization catalyst used may comprise a group VIII metal selected from Pt and Pd, alone or in combination and a support which may be amorphous or crystalline selected from alumina, amorphous silica alumina, ferrierite, ALPO-31, SAPO-11, SAPO-31, SAPO, 37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-11, MeAPO-31, MeAPO-41, MgAPSO-11, MgAPSO-31, MgAPSO-41, MgAPSO-46, ELAPSO-11, ELAPSO-3, EMAPSO-41, ELAPSO-11, ELAPSO-31, ELAPSO-41, laumontite, cancrinite, offertite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.

[0031] Patent application US 2017 / 0022424 A1 describes a process for hydrotreating renewable feedstocks for producing an n-paraffinic hydrocarbon fraction. This patent claims a very specific process implementation of injecting renewable feedstock into different catalytic beds, mixed with a gas recycle stream (rich in hydrogen) and / or with a liquid recycle stream. The recycled streams (gas and liquid) are obtained from an effluent separation section positioned downstream of the hydrotreatment stage and are injected according to a ratio relative to the renewable feedstock advantageously selected to control the reaction exotherm and the temperature profile of the reactor.A portion of the liquid stream comprising the paraffins is sent to a hydroisomerization step in the presence of a hydroisomerization catalyst comprising at least one metal from group VIII chosen from platinum and palladium, nickel and cobalt and / or at least one metal from group VI B chosen from molybdenum and tungsten and at least one amorphous support of doped alumina, silica alumina type or a zeolitic support preferably comprising a zeolite of structural type TON (Nll-10), FER (ferrierite), EUO (EU-1 or ZSM-50) or the zeolites ZSM-48, ZBM-30, IZM-1, COK-7, EU-2 and Ell-11 alone or as a mixture. All or part of the hydroisomerized effluent is then sent to a fractionation stage to obtain a gaseous fraction, a naphtha fraction and a middle distillate fraction comprising diesel and kerosene. Part of the middle distillate fraction (150°C. +) can be recycled in the hydrotreatment step or in the hydroisomerization step. In another variant, part of the 300°C fraction + can be recycled in the hydroisomerization stage so as to upgrade this cut into lighter products and to increase the cold properties. Unlike the process according to the invention, the heavy cut of the effluent from the hydroisomerization stage is not sent to a second hydroisomerization stage separate from the first.

[0032] Patent application US 2011 / 0105812 A1 describes a process for improving the cold properties of a diesel cut produced from renewable feedstocks. The process comprises a step of hydrogenation / deoxygenation of the renewable feedstocks allowing the production of an n-paraffinic effluent followed by a step of hydroisomerization and selective hydrocracking of the n-paraffinic effluent. The effluent collected after hydroisomerization and selective hydrocracking is then selectively separated to produce at least one diesel stream comprising a portion of the paraffins having a boiling point in the diesel range, at least one recycle stream comprising another portion of the paraffins having a boiling point in the diesel range (preferably C16+), the composition of the diesel stream being different from the composition of the recycle stream, at least one light stream containing at least LPGs and at least one naphtha-type stream.The recycle stream is advantageously recycled to the hydroisomerization and selective hydrocracking stage. The hydroisomerization catalyst used may comprise a group VIII metal selected from Pt and Pd, alone or in combination and a support which may be amorphous or crystalline selected from alumina, amorphous silica alumina, ferrierite, ALPO-31, SAPO-11, SAPO-31, SAPO, 37, SAPO-41, SM-3, MgAPSO-31, FU-9, NU-10, NU-23, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-50, ZSM-57, MeAPO-11, MeAPO-31, MeAPO-41, MgAPSO-11, MgAPSO-31, MgAPSO-41, MgAPSO-46, ELAPSO-11, ELAPSO-3, EMAPSO-41, ELAPSO-11, ELAPSO-31, ELAPSO-41, laumontite, cancrinite, offertite, stillbite in hydrogen form, mordenite in magnesium or calcium form, and partheite in calcium or magnesium form, alone or in combination.Unlike the process according to the invention, the recycle stream of the effluent from the hydroisomerization stage is not sent to a second hydroisomerization stage separate from the first.

[0033] Patent application WO 2021 / 099343 A1 describes the composition of a hydrocarbon fraction rich in iso-paraffins and more specifically the composition of a hydrocarbon fraction having an excellent freezing point, compatible with an aviation fuel application. More particularly, the object of the invention is to provide an aviation fuel type hydrocarbon fraction with increased yield and density and produced from a renewable feedstock. Said composition is obtained by a process comprising the hydrodeoxygenation of a renewable feedstock of vegetable oil type followed by a step of isomerization of the n-paraffins formed to produce iso-paraffins, the isomerization catalyst comprising a group VIII metal chosen from platinum, palladium, and nickel and a support comprising a zeolite chosen from SAPO-11, SAPO-41, ZSM-22, ZSM-23 or ferrierite.The hydroisomerized effluent obtained is then fractionated so as to obtain said composition. Unlike the process according to the invention, no cut resulting from the fractionation of the hydroisomerized effluent is sent to a second hydroisomerization stage separate from the first.

[0034] For the purposes of the present invention, the various embodiments presented can be used alone or in combination with each other, without limitation of combination.

[0035] For the purposes of the present invention, the different parameter ranges for a given step such as pressure ranges and temperature ranges may be used alone or in combination. For example, for the purposes of the present invention, a preferred range of pressure values ​​may be combined with a more preferred range of temperature values.

[0036] In the following text, the term naphtha cut or naphtha means the hydrocarbon fraction having a boiling point lower than the middle distillate cut. The middle distillate cut generally has an initial cutting point between 120 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 36°C boiling point) up to 216°C and includes the gasoline cut.

[0037] Throughout the rest of the text, the term kerosene or kerosene fraction is understood to mean the fraction having initial and final boiling points between 120 and 300°C and the term diesel or diesel fraction is understood to mean a fraction having initial and final boiling points in the range from 120 to 400°C and preferably between 120 and 380°C. In the rest of the text, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals in columns 8, 9 and 10 according to the new IIIPAC classification, and group VI B to the metals in column 6.

[0038] In the remainder of the text, the expressions "between ... and ..." and "between .... and ..." are equivalent and mean that the limit values ​​of the interval are included in the range of values ​​described. If this were not the case and the limit values ​​were not included in the range described, such precision will be provided by the present invention.

[0039] In this description, the expression "greater than..." is understood as strictly greater than, and symbolized by the sign ">", and the expression "less than" as strictly less than, and symbolized by the sign "<".

[0040] ADVANTAGES OF THE METHOD ACCORDING TO THE INVENTION

[0041] In attempting to develop a process for treating a feedstock from a renewable source for the selective production of a kerosene cut, the applicant discovered that the use of a specific bifunctional catalyst, comprising a hydrogenating / dehydrogenating phase based on a noble metal (Pt) and a support comprising a silica-alumina or comprising one or more specific zeolites chosen from zeolites with structural code MTW and IZM-2, combined with an implementation of the specific process including two hydroconversion stages, was of great interest.

[0042] Advantageously, the process according to the invention comprises three conversion stages, including a stage of hydrodeoxygenation of renewable feedstocks producing an effluent rich in linear paraffins and two separate stages of hydroconversion of the paraffinic effluent (hydroisomerization and / or hydrocracking reactions), the second hydroconversion stage using a specific catalyst.

[0043] Subject of the invention

[0044] More specifically, the present invention relates to a process for treating a feedstock from a renewable source to produce a kerosene cut comprising at least the following steps and preferably consisting of: a) a step of hydrotreating said feedstock in the presence of a fixed-bed catalyst, said catalyst comprising a hydrogenating function and an oxide support, at a temperature of between 200 and 450°C, at a pressure of between 1 and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 2000 Nm 3 hydrogen / m 3of charge, b) a step of separating at least a portion of the effluent from step a) into at least one light gaseous fraction, at least one hydrocarbon liquid effluent, and at least one aqueous liquid effluent, c) a first step of hydroconversion of at least a portion of the hydrocarbon liquid effluent from step b) in the presence of at least one bifunctional fixed-bed hydroconversion catalyst, said catalyst comprising a hydrogenating phase containing at least one metal from group VIII and / or at least one metal from group VIB of the periodic table, an acid 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 of between 250 and 500°C, at a pressure of between 1 and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3of charge, d) a step of separating at least a portion of the effluent from step c) which makes it possible to separate at least one gaseous fraction, and at least one liquid hydrocarbon effluent, e) a step of fractionating the liquid hydrocarbon effluent from step d) into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and a heavy cut having an initial boiling point of between 250 and 300°C, or a step of stabilizing the liquid hydrocarbon effluent from step d) allowing the separation of at least one light gaseous fraction, a naphtha hydrocarbon cut and a heavy cut having an initial boiling point of between 100 and 180°C,f) a second hydroconversion step distinct from step c) of all or part of the heavy cut having an initial boiling point of between 250 and 300°C from step e) in the case where step e) is a fractionation step or of all or part of the heavy cut having an initial boiling point of between 100 and 180°C from step e) in the case where step e) is a stabilization step, said second hydroconversion step being carried out in a fixed bed in the presence of a bifunctional hydroconversion catalyst comprising at least one noble metal, said noble metal being platinum and a support comprising a silica-alumina or of a bifunctional hydroconversion catalyst comprising at least one noble metal, said noble metal being platinum and a support comprising a zeolite chosen from zeolites with structural code MTW and IZM-2, alone or as a mixture,said hydroconversion step being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at an hourly space 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 Nm, 3 / m 3of feedstock treated in said step f), to produce a hydroconverted effluent, g) a step of separating at least a portion of the hydroconverted effluent from step f) which makes it possible to separate at least one gaseous fraction, and at least one liquid hydrocarbon effluent, h) a step of fractionating the liquid hydrocarbon effluent from step g) into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and a heavy cut having an initial boiling point of between 250 and 300°C, or a step of stabilizing the liquid hydrocarbon effluent from step g) allowing the separation of at least one light gaseous fraction, a naphtha hydrocarbon cut and a kerosene cut having an initial boiling point of between 100 and 180°C, said steps d) and g) and e) and h) being able to be carried out in the same equipment or in separate equipment.

[0045] An advantage of the present invention is to provide a process for treating a feedstock from a renewable source to produce one or more kerosene cuts, using a specific catalyst in a second hydroconversion step of all or part of the heavy cut from the first hydroconversion step, making it possible, for the cold property targets sought, to maximize the overall yield of said kerosene cut(s) of interest produced in said process.

[0046] Another advantage of the present invention is to provide the flexibility to co-produce a diesel cut at the outlet of the first hydroconversion step, said diesel cut having an initial boiling point between 100 and 180°C in the case where step e) is a stabilization or between 250 and 300°C in the case where step e) is a fractionation, with a kerosene cut at the outlet of the second hydroconversion step in the case where step e) is a stabilization or two kerosene cuts, one of which at the outlet of the first hydroconversion step and one at the outlet of the second hydroconversion step, in the case where step e) is a fractionation.

[0047] Another advantage of the present invention is to allow the treatment of heavier renewable source feedstocks within said process thanks to the use of a specific catalyst promoting the cracking of long paraffins in the first hydroconversion stage followed by a second hydroconversion stage which is more selective in hydroisomerization thanks to the use of another specific catalyst. Detailed description of the invention

[0048] Charges

[0049] The present invention is particularly dedicated to the preparation of diesel and / or kerosene fuel bases corresponding to new environmental standards, from charges from renewable sources.

[0050] The feedstocks from renewable sources used in the process according to the present invention are advantageously chosen from vegetable oils (for example palm, rapeseed, soybean), animal fats, used cooking oils, oils of microbial origin (for example from algae), fish oils, pine oils, long paraffins (waxes) from the Fischer-Tropsch process, crude or having undergone a pretreatment, or mixtures of such feedstocks, containing triglycerides and / or free fatty acids and / or esters. The vegetable oils may advantageously be crude or refined, totally or in part, and derived from the following plants: rapeseed, sunflower, soybean, palm, palm kernel, olive, coconut, jatropha, this list not being exhaustive. Algal or fish oils are also relevant.Animal fats are advantageously chosen from lard or fats composed of residues from the food industry or from the catering industry.

[0051] These fillers essentially contain triglyceride-type chemical structures that the skilled person also knows as fatty acid triester as well as free fatty acids. A fatty acid triester is thus composed of three fatty acid chains. These fatty acid chains in the form of triester or in the form of free fatty acids, have a number of unsaturations per chain, also called the number of carbon-carbon double bonds per chain, generally between 0 and 3 but which can be higher, particularly for oils derived from algae which generally have a number of unsaturations per chain of 5 to 6.

[0052] The molecules present in the fillers from renewable sources used in the present invention therefore have a number of unsaturations, expressed per triglyceride molecule, advantageously between 0 and 18. In these fillers, the unsaturation rate, expressed as the number of unsaturations per hydrocarbon fatty chain, is advantageously between 0 and 6. The fillers from renewable sources generally also contain various impurities and in particular heteroatoms such as nitrogen. The nitrogen contents in vegetable oils are generally between 1 and 1000 ppm by weight approximately, depending on their nature.

[0053] Process and catalysts

[0054] Step a)

[0055] Advantageously, the feedstock may undergo, prior to step a) of the process according to the invention, a pre-treatment or pre-refining step so as to eliminate, by appropriate treatment, contaminants such as metals, such as alkaline compounds, for example on ion exchange resins, alkaline earths and phosphorus. Suitable treatments may, for example, be thermal and / or chemical treatments well known to those skilled in the art.

[0056] In accordance with step a) of the process according to the invention, the feedstock, optionally pretreated, is brought into contact with a fixed-bed catalyst at a temperature of between 200 and 450°C, preferably between 220 and 350°C, preferably between 220 and 320°C, and even more preferably between 220 and 310°C. The pressure is between 1 and 10 MPa, preferably between 1 and 6 MPa and even more preferably between 1 and 4 MPa. The hourly space velocity, i.e. the volume of feedstock per volume of catalyst per hour, is between 0.1 and 10 h' 1 The charge is brought into contact with the catalyst in the presence of hydrogen. The total quantity of hydrogen mixed with the charge is such that the hydrogen / charge ratio is between 70 and 2000 Nm 3 hydrogen / m 3 load and preferably between 150 and 1000 Nm 3 hydrogen / m 3 dump.

[0057] In step a) of the process according to the invention, the fixed-bed catalyst is advantageously a hydrotreatment catalyst comprising a hydro-dehydrogenating function comprising at least one metal from group VIII and / or group VI B, taken alone or as a mixture and a support chosen from the group formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also advantageously contain other compounds and for example oxides chosen from the group formed by boron oxide, zirconia, titanium oxide, phosphoric anhydride. The preferred support is an alumina support and very preferably r], 5 or y alumina.

[0058] Said catalyst is advantageously a catalyst comprising metals from group VIII preferably chosen from nickel and cobalt, taken alone or in a mixture, preferably in association with at least one metal from group VI B preferably chosen from molybdenum and tungsten, taken alone or in a mixture.

[0059] The content of metal oxides of groups VIII and preferably nickel oxide is advantageously between 0.5% and 10% by weight of nickel oxide (NiO) and preferably between 1% and 5% by weight of nickel oxide and the content of metal oxides of groups VI B and preferably molybdenum trioxide is advantageously between 1% and 30% by weight of molybdenum trioxide (MoOa), preferably from 5% to 25% by weight, the percentages being expressed in % by weight relative to the total mass of the catalyst.

[0060] The total content of metal oxides from groups VI, B and VIII in the catalyst used in step a) is advantageously between 5% and 40% by weight and preferably between 6% and 30% by weight relative to the total mass of the catalyst.

[0061] Said catalyst used in step a) of the process according to the invention must advantageously be characterized by a high hydrogenating power so as to orient the selectivity of the reaction as much as possible towards a hydrogenation conserving the number of carbon atoms of the fatty chains, i.e. the hydrodeoxygenation route, in order to maximize the yield of hydrocarbons entering the distillation field of kerosenes and / or diesels. This is why it is preferred to operate at a relatively low temperature. Maximizing the hydrogenating function also makes it possible to limit the polymerization and / or condensation reactions leading to the formation of coke which would degrade the stability of the catalytic performances. Preferably, a Ni or NiMo type catalyst is used.

[0062] Said catalyst used in hydrotreatment step a) of the process according to the invention may also advantageously contain a doping element chosen from phosphorus and boron, taken alone or in a mixture. Said doping element may be introduced into the matrix or preferably be deposited on the support. Silicon may also be deposited on the support, alone or with phosphorus and / or boron and / or fluorine.

[0063] The weight content of oxide of said doping element is advantageously less than 20% by weight and preferably less than 10% by weight and it is advantageously at least 0.001% by weight.

[0064] Preferred catalysts are the catalysts described in patent application FR 2 943 071 describing catalysts having high selectivity for hydrodeoxygenation reactions. Other preferred catalysts are the catalysts described in patent application EP 2 210 663 describing supported or bulk catalysts comprising an active phase consisting of a sulfurized element from group VIB, the element from group VIB being molybdenum.

[0065] The metals of the catalysts used in step a) of hydrotreatment of the process according to the invention are sulphide metals or metallic phases and preferably sulphide metals.

[0066] It would not be outside the scope of the present invention to use in step a) of the process according to the invention, simultaneously or successively, a single catalyst or several different catalysts. This step can be carried out industrially in one or more reactors with one or more catalytic beds and preferably with a descending liquid flow.

[0067] Said hydrotreatment step a) allows the hydrogenation, hydrodeoxygenation, hydrodenitrogenation and hydrodesulfurization of said feed.

[0068] Step b)

[0069] In accordance with step b) of the process according to the invention, a step of separating at least a portion and preferably all of the effluent from step a) is carried out. Said step b) makes it possible to separate at least one so-called light gaseous fraction rich in hydrogen, at least one liquid hydrocarbon effluent consisting of n-paraffins, and at least one aqueous liquid effluent.

[0070] Said light gaseous fraction comprises at least the hydrogen not converted by the reactions carried out in step a), at least the gases with one or more oxygen atoms resulting from the decomposition of the oxygenated compounds in step a) and at least the C4 compounds; that is to say the compounds C1 to C4 preferably having a final boiling point of less than 20°C. The aim of this step is to separate the gases from the liquids. More particularly, the aim is to recover at least the hydrogen-rich gases which may also contain compounds such as CO and CO2, at least one liquid hydrocarbon effluent consisting of n-paraffins and at least one aqueous liquid effluent containing the water produced by the reactions carried out in step a). Said liquid hydrocarbon effluent preferably has a sulfur content of less than 10 ppm by weight, a nitrogen content of less than 2 ppm by weight.

[0071] Separation step b) can advantageously be implemented by any method known to those skilled in the art, such as, for example, the combination of one or more high and / or low pressure separators operated hot or cold, and / or high pressure and / or low pressure stripping.

[0072] Step b) also allows the separation of at least one aqueous liquid effluent, preferably water. The removal of at least part of the water and preferably all of the water can be carried out by any methods and techniques known to those skilled in the art. Preferably, the removal of the water is carried out by decantation in a separator drum or by drying or by passing through a desiccant or by flash or by a combination of at least two of these techniques. The atomic oxygen content of the liquid hydrocarbon effluent containing the paraffinic hydrocarbons resulting from step b) of the process according to the invention, expressed in parts per million by weight (ppm), is preferably less than 10,000 ppm, preferably less than 6,000 ppm, very preferably less than 1,000 ppm by weight, even more preferably less than 500 ppm by weight.The content in ppm by weight of atomic oxygen in said hydrocarbon liquid effluent is measured by the infrared absorption technique such as for example the technique described in patent application US2009 / 0018374A1.

[0073] Step c)

[0074] According to the invention, the process comprises a first step c) of hydroconversion of at least a portion and preferably all of the hydrocarbon liquid effluent from step b) of the process in the presence of a fixed-bed bifunctional 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 as a mixture, an acid 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 of between 250 and 500°C, at a pressure of between 1 and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3load, and preferably between 150 and 750 Nm 3 / m 3 dump.

[0075] The operating conditions of the first hydroconversion step c) are adjusted to promote the hydroisomerization or hydrocracking reactions as required. Preferably, the hydroconversion step c) advantageously operates at a temperature of between 250 and 450°C, and very preferably between 250 and 400°C, at a pressure of between 2 and 10 MPa and very preferably between 3 and 9 MPa, at an hourly volumetric flow rate advantageously of between 0.2 and 7 h -1 and very preferably, between 0.5 and 5 h -1 at a hydrogen flow rate such that the hydrogen / charge volume ratio is advantageously between 100 and 1000 Nm 3 / m 3 load and preferably between 150 and 1000 Nm 3 / m 3 dump.

[0076] Said first hydroconversion step may advantageously comprise one or more catalytic beds which may comprise one or more different catalysts in each of the beds.

[0077] 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 VI B of the periodic table alone or as a mixture, an acid support comprising a silica-alumina or one or more zeolites and optionally at least one binder.

[0078] The hydro / dehydrogenating function

[0079] The metals of group VIII are advantageously chosen from iron, cobalt, nickel, platinum, and palladium, taken alone or in a mixture, and preferably from nickel, cobalt, platinum, and palladium.

[0080] Group VI B metals are chosen from tungsten and molybdenum, taken alone or in mixture.

[0081] In the case where the metals of group VIII are chosen from non-noble metals, the following combinations of metals are preferred: nickel-molybdenum, cobalt-molybdenum, nickel-tungsten, cobalt-tungsten, and very preferably: nickel-molybdenum, nickel-tungsten. It is also possible to use combinations of three metals such as for example nickel-cobalt-molybdenum.

[0082] The content of the non-noble group VIII metal catalyst 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 very preferably between 1% and 4% by weight of oxide.

[0083] The content of the catalyst in metal from group VI B is advantageously between 1% and 30% by weight of oxide relative to the total weight of said catalyst, preferably between 2% and 25% by weight of oxide, very preferably between 5% and 20% by weight of oxide, and even more preferably between 5% and 16% by weight of oxide.

[0084] The catalyst used in step c) may also advantageously further comprise at least one additional metal chosen from the group formed by the metals of groups II IA, IVA and VI IB of the periodic table of elements and preferably chosen from gallium, indium, tin and rhenium. Said additional metal is preferably chosen from indium, tin and rhenium.

[0085] 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.

[0086] According to one embodiment, the sulfur content in the hydroconversion catalyst is such that the ratio of the number of moles of sulfur to the number of moles of the at least one noble group VIII metal is between 0.3 and 3. According to one or more embodiments, the presence of sulfur in the catalyst comes from an optional sulfurization step of the hydroconversion catalyst. According to one or more embodiments, the presence of sulfur in the catalyst comes from potentially present impurities, such as for example in the alumina binder. According to another embodiment, the catalyst does not contain sulfur.

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

[0088] In a preferred embodiment, the catalyst comprises at least one metal from group VIII and preferably nickel and at least one metal from group VI B and preferably tungsten, preferably active in their sulfurized form.

[0089] In another embodiment, the catalyst comprises at least one group VIII metal and preferably a group VIII noble metal chosen from platinum and palladium. Preferably, the group VIII metal of the catalyst used in step c) is platinum, preferably active in its reduced form.

[0090] The metal function is advantageously introduced into the catalyst by any method known to those skilled in the art, such as for example co-mixing, dry impregnation or exchange impregnation.

[0091] The acid function.

[0092] According to the invention, the catalyst used in step c) comprises at least one acid support comprising a silica-alumina or one or more zeolites. In the case where the acid support consists of one or more zeolites, said zeolites are chosen from zeolites of structural type FAU, *BEA, ISV, IWR, IWW, MEI, UWY, taken alone or as a mixture and preferably chosen from zeolites of structural type FAU and *BEA, taken alone or as a mixture. In a preferred embodiment, the zeolite is chosen from zeolite Y and zeolite beta taken alone or as a mixture and preferably the zeolite is zeolite Y and very preferably dealuminated zeolite USY.

[0093] 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 more preferably between 4 and 20% by weight relative to the total weight of said catalyst.

[0094] Preferably, the acidic support of the catalyst used in step c) comprises and preferably consists of silica-alumina.

[0095] The binder

[0096] Optionally, said catalyst support used in step c) may also comprise a binder. Preferably, said support comprises a binder when said support comprises a zeolite. Said binder is advantageously chosen from silica (SiCh), alumina (AI2O3), clays, titanium oxide (TiCh), boron oxide (B2O3) and zirconia (ZrCh) taken alone or as a mixture. 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.

[0097] A preferred hydroconversion catalyst used in first step c) of hydroconversion comprises nickel and tungsten, and a silica-alumina as an acidic support, without any other binder. In this case, said catalyst is in the sulfurized form.

[0098] A preferred catalyst comprises a silica-alumina and at least tungsten and / or molybdenum and at least nickel and / or cobalt, said catalyst being sulfurized. The content of tungsten and / or molybdenum is advantageously comprised, in oxide equivalent, between 5% and 50% by weight relative to the finished catalyst, preferably between 10% and 40% by weight and very preferably between 15% and 35% by weight and the content of nickel and / or cobalt of said catalyst is advantageously comprised, in oxide equivalent, between 0.5% and 10% by weight relative to the finished catalyst, preferably between 1% and 8% by weight and very preferably between 1.5% and 6% by weight. The element content is perfectly measured using X-ray fluorescence. A preferred catalyst used in the first step c) of hydroconversion, comprises a particular silica-alumina, said silica-alumina having:

[0099] - alumina and silica with a mass content of silica (SiCh) 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 is advantageously between 10 and 50% by weight

[0100] - a BET specific surface area of ​​100 to 500 m 2 / g, preferably between 200 m 2 / g and 450 m 2 / g and very preferably between 200 m 2 / g and 300 m 2 / g,

[0101] - an average diameter of the mesopores measured by mercury porosimetry of between 3 and 12 nm, preferably between 3 nm and 11 nm and very preferably between 4 nm and 10.5 nm,

[0102] - 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 very preferably between 0.4 and 0.8 ml / g,

[0103] - a volume of macropores, whose diameter is greater than 50 nm, less than 0.002 ml / g. 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 μm.

[0104] Preferably, the metal distribution coefficient of said preferred catalyst is greater than 0.1, preferably greater than 0.2 and very 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 Castaing microprobe.

[0105] Another preferred hydroconversion catalyst used in first hydroconversion step c) comprises at least one noble metal, said noble metal being platinum, and a silica-alumina as an acidic support, without any other binder. In this case, said catalyst is in reduced form.

[0106] In this case, the silica content of the silica-alumina, expressed as a weight percentage, is advantageously between 1% and 95%, advantageously between 5% and 95% and preferably between 10% and 80% and very preferably between 20% and 70% and even more preferably between 22% and 45%. This silica content is perfectly measured using X-ray fluorescence.

[0107] A preferred hydroconversion catalyst used in first step c) of hydroconversion comprises a particular silica-alumina. Preferably, said catalyst comprises 0.05% to 10% by weight, preferably between 0.1% and 5% by weight of at least one noble metal from group VIII, preferably chosen from platinum and palladium (preferably platinum) deposited on the silica-alumina, without any other binder, containing a quantity of silica (SiC>2) of between 1% and 95%, expressed as a weight percentage, preferably between 5% and 95%, preferably between 10% and 80% and very preferably between 20% and 70% and even more preferably between 22% and 45%, said catalyst having: a BET specific surface area of ​​100 to 500 m 2 / g, preferably between 200 and 450 m 2 / g and very preferably between 200 and 300 m 2 / g, an average diameter of the mesopores measured by mercury porosimetry of between 4 and 12 nm, preferably of between 4 and 11 nm and very preferably of between 5 and 11 nm, a total pore volume measured by mercury porosimetry of between 0.2 and 1.2 ml / g, preferably of between 0.3 and 1.0 ml / g and very preferably of between 0.3 and 0.9 ml / g, a volume of the macropores measured by mercury porosimetry, the diameter of which is greater than 50 nm, less than 0.02 ml / g. a content of alkali or alkaline-earth compounds of less than 300 ppm by weight and preferably less than 200 ppm by weight.

[0108] The average mesopore diameter is defined as the diameter corresponding to the cancellation of the curve derived from the mercury intrusion volume obtained by mercury porosimetry for pore diameters between 3.7 and 50 nm.

[0109] Preferably, the dispersion of the noble metal of said preferred catalyst is advantageously between 5% and 100%, preferably between 5% and 90% and very preferably between 10% and 90%. The dispersion, representing the fraction of metal accessible to the reagent relative to the total quantity of metal of the catalyst, is advantageously measured, for example, by H2 / O2 titration or by transmission electron microscopy.

[0110] Preferably, the noble metal distribution coefficient of said preferred catalyst is greater than 0.1, preferably greater than 0.2 and very preferably greater than 0.4. The noble metal distribution represents the distribution of the metal within the catalyst grain, the metal being able to be well or poorly dispersed. Thus, it is possible to obtain poorly distributed platinum (for example detected in a crown whose thickness is significantly less than the radius of the grain), but well dispersed, that is to say that all the platinum atoms, located in the crown, will be accessible to the reagents. The noble metal distribution coefficient can be measured by Castaing microprobe. Preferably, the catalyst used in the first step c) is shaped in the form of cylindrical or polylobed extrudates such as bilobed, trilobed, polylobed of straight or twisted shape.According to one or more embodiments, the catalyst used in step c) is shaped into crushed powders, tablets, rings, balls, wheels. Techniques other than extrusion, such as pelletizing or coating, can advantageously be used.

[0111] Advantageously, the catalyst of the first hydroconversion stage may be bifunctional in nature with a hydrogenating / dehydrogenating phase of the nickel-tungsten oxide or reduced platinum type on a silica-alumina acid support. This specific choice advantageously makes it possible to co-process atypical feedstocks in a mixture with renewable feedstocks, for example waxes from the Fischer-Tropsch process known to be rich in heavy paraffins, and to maximize their conversion by hydrocracking to the middle distillate distillation range. Sending the converted n-paraffins obtained to a second hydroconversion stage operating with a specific Pt-type catalyst on a crystalline support promoting hydroisomerization also makes it possible to maximize the conversion to the kerosene cut.

[0112] According to the invention, the process comprises a step d) of separating at least part and preferably all of the effluent from step c).

[0113] Step d)

[0114] Said step d) makes it possible to separate at least one so-called light gaseous fraction and at least one hydrocarbon liquid effluent. Optionally, step d) also makes it possible to separate at least part of the residual water and preferably all of the residual water.

[0115] Said so-called light gaseous fraction comprises at least the hydrogen not converted by the reactions described in step c) and at least a portion of the C4 cracking products; i.e. compounds C1 to C4 preferably having a final boiling point below 20°C. The purpose of this step is to separate the gases from the liquids. More particularly, the purpose is to recover at least the hydrogen-rich gases, at least one liquid hydrocarbon effluent rich in branched paraffins and optionally an aqueous liquid effluent containing a small amount of residual water (residual water generated either by the hydrodeoxygenation of the residual oxygenated compounds present in the feedstock of step c) or by a slight entrainment of water from step b) to step c)).

[0116] The separation step d) described can advantageously be implemented by any method known to those skilled in the art such as, for example, the combination of one or more high and / or low pressure separator tanks operated hot or cold, and / or high pressure and / or low pressure stripping.

[0117] Step e)

[0118] In accordance with the process according to the invention, the liquid hydrocarbon effluent from step d) undergoes a step e) of fractionation into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and preferably between 110 and 130°C and a heavy cut having an initial boiling point of between 250 and 300°C and preferably between 270 and 300°C or a stabilization step e) allowing the separation of at least one light gaseous fraction, a naphtha hydrocarbon cut and a heavy cut having an initial boiling point of between 100 and 180°C, and preferably between 110 and 130°C.

[0119] Fractionation step e) can advantageously be carried out by any method known to those skilled in the art and can preferably be carried out in a distillation column or in a steam stripping step followed by a distillation column.

[0120] Said heavy cut having an initial boiling point between 250 and 300°C separated in said step e) is advantageously obtained at the bottom of the distillation column.

[0121] More particularly, the cut point between the kerosene and naphtha cut is adjustable in the range between 100 and 180°C, and preferably around 120°C. More particularly, the cut point between the kerosene cut and the heavy cut is adjustable in the range between 250 and 300°C if the aim is to target the production of kerosene, and preferably around 280°C.

[0122] The naphtha cut thus contains compounds which have a boiling point below 120°C, the kerosene cut contains compounds which have a boiling point between 120 and 280°C.

[0123] The recovery of the bio-naphtha cut is not the subject of the present invention, but this cut can advantageously be sent to a steam cracking or reforming unit or even recovered as a gasoline base in a mixture with other gasoline bases.

[0124] In the case where step e) is a stabilization step, said step e) allows the separation of three hydrocarbon cuts: a light gaseous fraction, a naphtha hydrocarbon cut and a heavy cut having an initial boiling point of between 100 and 180°C, and preferably between 110 and 130°C. Stabilization step e) can be implemented by any method known to those skilled in the art such as, for example, a gas stripping step, advantageously with steam, and / or separation within separation tanks and / or a combination of these options.

[0125] Preferably, said stabilization step is a step of stripping at least part and preferably all of the liquid hydrocarbon effluent from step d).

[0126] Optionally, the heavy cut obtained at the end of step e) is sent in whole or in part only to the second hydroconversion step f). In the case where a part of the heavy cut from step e) is not directed to step f), said process advantageously makes it possible to co-produce a bio-diesel cut with the required fuel properties and thus offers the flexibility to the operator of said process to produce various bio-fuel bases on the one hand (road bio-diesel and bio-kerosene), and to adjust their respective yields on the other hand by adapting the proportion of the heavy cut directed to the second hydroconversion step f).

[0127] Step f)

[0128] According to the invention, said process comprises a second step f) of hydroconversion, distinct from step c), of all or part and preferably all of the heavy cut having an initial boiling point of between 250 and 300°C from step e) in the case where said step e) is a fractionation step or of all or part and preferably all of the heavy cut having an initial boiling point of between 100°C and 180°C from step e) in the case where step e) is a stabilization step.

[0129] Optionally according to the invention, a second external paraffinic feedstock may advantageously be sent to said second hydroconversion step f). Said second external paraffinic feedstock to said process according to the invention is advantageously sent to said step f) in a mixture with said heavy cut. For example, this second feedstock may come from a hydrodeoxygenation step of a renewable feedstock distinct from step a) described in said process. The origin described here is not limiting and other origins of external feedstocks are conceivable.

[0130] According to the invention, said hydroconversion step f) is carried out in the presence of a fixed-bed bifunctional hydroconversion catalyst comprising at least one noble metal, said noble metal being platinum, and a support comprising a silica-alumina, or a fixed-bed bifunctional hydroconversion catalyst comprising at least one noble metal, said noble metal being platinum, and a support comprising a zeolite chosen from zeolites with structural code MTW and IZM-2, alone or as a mixture, said hydroconversion step being carried out at a temperature of between 250 and 500°C, at a pressure of between 1 and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3 of total charge treated in said step f) to produce a hydroconverted effluent.

[0131] The operating conditions of the second hydroconversion step f) are adjusted to promote the hydroisomerization or hydrocracking reactions as required. Preferably, the second hydroconversion step f) advantageously operates at a temperature of between 250 and 450°C, and very preferably, between 250 and 400°C, at a pressure of between 2 and 10 MPa and very preferably, between 3 and 9 MPa, at an hourly volumetric flow rate advantageously of between 0.2 and 7 h -1 and very preferably, between 0.5 and 5 h-1, at a hydrogen flow rate such that the hydrogen / charge volume ratio is advantageously between 100 and 1000 normal m 3 of hydrogen per m 3 load and preferably between 150 and 1000 normal m 3 of hydrogen per m 3 of total charge treated in said step f).

[0132] According to the invention, the bifunctional catalyst used in said second hydroconversion step f) comprises at least one noble metal, said noble metal being platinum and a support comprising a silica-alumina or said catalyst comprises at least one noble metal, said noble metal being platinum and a support comprising a zeolite chosen from zeolites with structural code MTW and IZM-2, alone or as a mixture,

[0133] Preferably, the content of noble metal from group VIII, and preferably the platinum content, in the catalyst used in step f) is between 0.01% and 4% by weight, preferably between 0.05% and 2% by weight, relative to the total weight of said catalyst.

[0134] In the case where the bifunctional catalyst comprises a support comprising a silica-alumina, the silica content of the silica-alumina, expressed as a weight percentage, is advantageously between 1% and 95%, advantageously between 5% and 95% and preferably between 10% and 80% and very preferably between 20% and 70% and even more preferably between 22% and 45%. This silica content is perfectly measured using X-ray fluorescence.

[0135] A preferred hydroconversion catalyst used in step f) of the process according to the invention comprises a particular silica-alumina. Preferably, said catalyst comprises 0.05% to 10% by weight, preferably between 0.1% and 5% by weight of at least one noble metal from group VIII, preferably chosen from platinum and palladium (preferably platinum) deposited on a silica-alumina support, without any other binder, containing a quantity of silica (SiCh) of between 1% and 95%, expressed as a weight percentage, preferably between 5 and 95%, preferably between 10% and 80% and very preferably between 20% and 70% and even more preferably between 22% and 45%, said catalyst having: a BET specific surface area of ​​100 to 500 m 2 / g, preferably between 200 and 450 m 2 / g and very preferably between 200 and 300 m 2 / g, an average diameter of the mesopores measured by mercury porosimetry of between 4 and 12 nm, preferably of between 4 and 11 nm and very preferably of between 5 and 11 nm, a total pore volume measured by mercury porosimetry of between 0.2 and 1.2 ml / g, preferably of between 0.3 and 1.0 ml / g and very preferably of between 0.3 and 0.9 ml / g, a volume of the macropores measured by mercury porosimetry, the diameter of which is greater than 50 nm, less than 0.02 ml / g. a content of alkali or alkaline-earth compounds of less than 300 ppm by weight and preferably less than 200 ppm by weight.

[0136] The average mesopore diameter is defined as the diameter corresponding to the cancellation of the curve derived from the mercury intrusion volume obtained by mercury porosimetry for pore diameters between 3.7 and 50 nm.

[0137] Preferably, the dispersion of the metal of said preferred catalyst is advantageously between 5% and 100%, preferably between 5% and 90% and very preferably between 10% and 90%. The dispersion, representing the fraction of metal accessible to the reagent relative to the total quantity of metal of the catalyst, is advantageously measured, for example, by H2 / O2 titration or by transmission electron microscopy.

[0138] Preferably, the noble metal distribution coefficient of said preferred catalyst is greater than 0.1, preferably greater than 0.2 and very preferably greater than 0.4. The noble metal distribution represents the distribution of the metal within the catalyst grain, the metal being able to be well or poorly dispersed. Thus, it is possible to obtain poorly distributed platinum (for example detected in a crown whose thickness is significantly less than the radius of the grain), but well dispersed, that is to say that all the platinum atoms, located in the crown, will be accessible to the reagents. The noble metal distribution coefficient can be measured by Castaing microprobe. Another preferred catalyst of the second hydroconversion step f) comprises and is preferably constituted by platinum, and a support comprising and preferably constituted by a zeolite with structural code MTW, preferably ZSM-12, and an alumina binder.

[0139] Another preferred catalyst of the second stage f) of hydroconversion comprises and is preferably constituted by platinum, and a support comprising and preferably constituted by an IZM-2 zeolite and an alumina binder.

[0140] The zeolites are preferably essentially in acid form, that is to say that the atomic ratio between the monovalent compensation cation (for example sodium) and the aluminum inserted in the crystal lattice of the solid is advantageously less than 0.1, preferably less than 0.05 and very preferably less than 0.01. According to one or more embodiments, the zeolites entering into the composition of said hydroconversion catalyst are advantageously calcined. According to one or more embodiments, said zeolites are exchanged by at least one treatment with a solution of at least one ammonium salt so as to obtain the ammonium form of the zeolites which, once calcined, leads to the acid form of said zeolites.

[0141] According to a preferred embodiment, the catalyst used in step f) more particularly comprises, and preferably consists of:

[0142] - from 1% to 90% by weight, preferably from 3% to 80% by weight and even more preferably from 4% to 60% by weight of zeolite;

[0143] - from 0.01% to 4% by weight, preferably from 0.05% to 2% by weight of at least one metal from group VI II B, preferably platinum;

[0144] - optionally from 0.01% to 2% by weight, preferably from 0.05% to 1% by weight of at least one additional metal chosen from the group formed by the metals of groups IIIA, IVA and VII B;

[0145] - optionally a sulfur content, preferably such that the ratio of the number of moles of sulfur to the number of moles of metal(s) of group VI 11 B is between 0.3 and 3; and

[0146] - optionally at least one binder, preferably alumina, ensuring the 100% complement in the catalyst, relative to the total weight of the catalyst from step e).

[0147] The catalysts used in the first and second steps c) and f) may advantageously be identical or different and preferably different.

[0148] In a preferred embodiment, the use of said specific catalyst in the second step f) of hydroconversion of all or part of the heavy cut having an initial boiling point of between 250 and 300°C from step e) in the case where step e) is a fractionation step or of all or part of the heavy cut having an initial boiling point of between 100 and 140°C from step e) in the case where step e) is a stabilization step, optionally mixed with a second external paraffinic feedstock, makes it possible, for the cold property targets sought, to maximize the yield of said kerosene cut of interest produced in said process.

[0149] Step g)

[0150] According to the invention, the process comprises a step g) of separating at least part and preferably all of the effluent from step f).

[0151] Said step g) makes it possible to separate at least one so-called light gaseous fraction and at least one hydrocarbon liquid effluent. Optionally, step g) also makes it possible to separate at least part of the residual water and preferably all of the residual water.

[0152] Said so-called light gaseous fraction comprises at least the hydrogen not converted by the reactions described in step f) 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 optionally an aqueous liquid effluent containing a small amount of residual water (residual water generated either by the hydrodeoxygenation of the residual oxygenated compounds present in the feedstock of step f) or by a slight entrainment of water from step e) to step f)).

[0153] The separation step g) described can advantageously be implemented by any method known to those skilled in the art such as, for example, the combination of one or more high and / or low pressure separator tanks operated hot or cold, and / or high pressure and / or low pressure stripping.

[0154] Step h)

[0155] According to the invention, the process comprises a step h) of fractionating the liquid hydrocarbon effluent from step g) into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and a heavy cut having an initial boiling point of between 250 and 300°C, or a step of stabilizing the liquid hydrocarbon effluent from step g) allowing the separation of at least one light gaseous fraction, a naphtha hydrocarbon cut and a kerosene cut having an initial boiling point of between 100 and 180°C, said steps d) and g) and e) and h) being able to be carried out in the same equipment or in separate equipment.In the case where step h) is a stabilization step, the stabilization step h) can be implemented by any method known to those skilled in the art such as for example a stripping step advantageously using steam and / or separation within separation tanks and / or a combination of these options.

[0156] Preferably, said stabilization step is a step of stripping at least part and preferably all of the liquid hydrocarbon effluent from step g).

[0157] According to the invention, said stripping step allows the separation of a light gaseous fraction, a naphtha hydrocarbon cut and a kerosene cut.

[0158] More particularly, the cut point between the kerosene and naphtha cuts is adjustable in the range 100 to 180°C, preferably around 120°C. The naphtha cut thus contains the compounds which have a boiling point lower than 120°C and the kerosene cut contains the compounds which have a boiling point higher than 120°C. In the presence of the specific catalyst in accordance with the invention, the operating conditions of the second hydroconversion step f) can advantageously be adjusted to selectively produce a bio-kerosene cut.

[0159] In a variant, step h) is a step of fractionating the liquid hydrocarbon effluent from step g) into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and preferably between 110 and 130°C and a heavy cut having an initial boiling point of between 250 and 300°C and preferably between 270 and 300°C.

[0160] Fractionation step h) can advantageously be carried out by any method known to those skilled in the art and can preferably be carried out in a distillation column or in a steam stripping step followed by a distillation column.

[0161] Said heavy cut having an initial boiling point between 250 and 300°C separated in said step h) is advantageously obtained at the bottom of the distillation column.

[0162] More particularly, the cut point between the kerosene and naphtha cut is adjustable in the range between 100 and 180°C, and preferably around 120°C. More particularly, the cut point between the kerosene cut and the heavy cut is adjustable in the range between 250 and 300°C, and preferably around 280°C, if the aim is to target kerosene production.

[0163] The recovery of the bio-naphtha cut from step h) is not the subject of the present invention, but this cut can advantageously be sent to a steam cracking or reforming unit or even recovered as a gasoline base in a mixture with other gasoline bases.

[0164] According to the invention, said steps d) and g) and e) and h) can be carried out in the same equipment or in separate equipment. Thus, said steps d) and g) can be carried out in the same equipment or in separate equipment and said steps e) and h) can be carried out in the same equipment or in separate equipment.

[0165] In one embodiment, said second hydroconversion step f) as well as steps g) and h) can be carried out on the same site as the previous steps a) to e) of the process according to the invention or on a different site, i.e. ex-situ or relocated in a dedicated hydroconversion unit not belonging to the process according to the invention.

[0166] The process according to the invention advantageously does not include a hydrocracking step downstream of step c) of hydroconversion of part or all of the effluent from step c).

[0167] Description of figures

[0168] [Fig 1] Figure 1 represents the different stages of the renewable kerosene production process comprising a hydrodeoxygenation stage and two hydroconversion stages and using a specific catalyst for the second hydroconversion stage f).

[0169] The feedstock from renewable sources is sent via line 1 mixed with make-up and / or recycled hydrogen (2) to a hydrotreatment unit a). The hydrotreated effluent from hydrotreatment unit a) is withdrawn via line 3 and sent to three-phase separation unit b) which separates at least one hydrogen-rich gaseous effluent (4) and at least one hydrocarbon liquid effluent (5). Unit b) also removes at least part of the water produced by the hydrodeoxygenation reactions and preferably all of this water (6).

[0170] The liquid hydrocarbon effluent (5) is sent to a first hydroconversion unit c), in the presence of a make-up and / or recycled hydrogen flow (7) to produce a second effluent (8) which is sent to a three-phase separation unit d) making it possible to separate a hydrogen-rich gaseous effluent (9) which may also contain light products such as the C1 - C4 cut, and at least one liquid hydrocarbon effluent (10). Said unit d) also makes it possible to eliminate at least part of the water and preferably all of the residual water (11).

[0171] The liquid hydrocarbon effluent (10) from unit d) is sent to a stripping stabilization unit e) allowing the separation of a light gaseous fraction (12), a naphtha hydrocarbon cut (13), and a diesel cut (15) having an initial boiling point of between 100 and 180°C. Optionally, unit e) is a fractionation unit allowing the production of a first bio-kerosene cut (14) and a diesel cut having an initial boiling point of between 250 and 300°C (15) by distillation of the hydrocarbon effluent (10) from unit d).

[0172] The diesel cut (15) is then directed to a second hydroconversion unit f), in the presence of a make-up and / or recycled hydrogen flow (18) to produce a third effluent (19).

[0173] Optionally, part of the diesel cut is directed to a finished product storage unit (16) allowing the co-production of a bio-diesel road fuel base.

[0174] Optionally, a second external stream rich in n-paraffins (17) can be mixed with the diesel cut (15) from unit e) in order to be co-treated with the diesel cut (15) in the second hydroconversion unit f).

[0175] The third effluent (19) is sent to a three-phase separation unit g) for separating a hydrogen-rich gaseous effluent (20) which may also contain light products such as the C1 - C4 cut, and at least one hydrocarbon liquid effluent (21). Said unit g) also allows the elimination of at least part of the water and preferably all of the residual water (22).

[0176] The liquid hydrocarbon effluent (21) from unit g) is sent to a stripping stabilization unit h) allowing the separation of a light gaseous fraction (23), a naphtha hydrocarbon cut (24), and a bio-kerosene cut (25) having an initial boiling point between 100 and 180°C.

[0177] Optionally, unit h) is a fractionation of the liquid hydrocarbon effluent (21) from unit g) into a light gaseous fraction (23), a naphtha hydrocarbon cut (24), a bio-kerosene cut (25) having an initial boiling point of between 100 and 180°C and a bio-diesel hydrocarbon cut having an initial boiling point of between 250 and 300°C.

[0178] The examples below illustrate the invention without limiting its scope. EXAMPLES

[0179] Example 1: preparation of a hydrotreatment catalyst (C1).

[0180] The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with molybdenum oxide MoOa contents of 22% by weight, nickel oxide NiO of 4% by weight and phosphorus oxide P2O5 of 5% by weight relative to the total weight of the finished catalyst.

[0181] Example 2: Preparation of a hydroconversion catalyst in accordance with the invention (C2).

[0182] The silica-alumina powder is prepared according to the synthesis protocol described in patent EP1 415 712A. The quantities of orthosilicic acid and aluminum hydrate are chosen so as to have a composition of 70% by weight of alumina AI2O3 and 30% by weight of silica SiC>2 in the final solid.

[0183] 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 at the mill outlet is 8% relative to the silica-alumina mixed solid. Then the suspension is conventionally dried in an atomizer in a conventional manner from 300°C to 60°C. The powder thus prepared 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 equipped with orifices of diameter 1.4 mm. The extrudates thus obtained are dried in an oven at 140°C then calcined under a flow of dry air at 550°C then calcined at 850°C in the presence of water vapor.

[0184] The characteristics of the support thus prepared are as follows: an average diameter of the mesopores measured by mercury porosimetry of 7.7 nm, a total pore volume of 0.49 ml / g, a mesoporous volume of 0.47 ml / g, a volume of the macropores, the diameter of which is greater than 50 nm less than 0.01 ml / g, a BET surface area of ​​240 m 2 / g,

[0185] The silica-alumina extrudates are then subjected to a dry impregnation step with an aqueous solution of ammonium metatungstate and nickel nitrate left to mature in a water maturer for 24 hours at room temperature and then calcined for two hours in dry air in a crossed bed at 450°C (temperature rise ramp of 5°C / min). The weight content of tungsten oxide WO3 of the finished catalyst after calcination is 27%, the nickel oxide NiO content is 3.5%. The metal distribution coefficient measured by Castaing microprobe is equal to 0.93.

[0186] Example 3: Preparation of a hydroconversion catalyst in accordance with the invention (C3).

[0187] Synthesis of IZM-2 zeolite.

[0188] The IZM-2 zeolite was synthesized in accordance with the teaching of patent FR 2 918 050 B. A colloidal suspension of silica known under the commercial term Ludox HS-40 marketed by Aldrich, is incorporated into a solution composed of sodium hydroxide (Prolabo), structuring agent dibromide of 1,6bis(methylpiperidinium)hexane, aluminum hydroxide (Aldrich) and deionized water. The molar composition of the mixture is as follows: 1 SiCh; 0.0060 AI2O3; 0.1666 Na2O; 0.1666 1,6bis(methylpiperidinium)hexane; 33.3333 H2O. The mixture is stirred vigorously for half an hour. The mixture is then transferred, after homogenization, into a PARR type autoclave. The autoclave is heated for 5 days at 170°C with stirring on a rotating spit (30 rpm). The product obtained is filtered, washed with deionized water to reach a neutral pH and then dried overnight at 100°C in an oven. The solid is then introduced into a muffle furnace to be calcined to remove the structuring agent.The calcination cycle includes a temperature rise to 200°C, a two-hour hold at this temperature, a temperature rise to 550°C followed by an eight-hour hold at this temperature and finally a return to room temperature. The temperature rises are carried out with a ramp of 2°C / min. The solid thus obtained is then refluxed for 2 hours in an aqueous solution of ammonium nitrate (10 ml of solution per gram of solid, ammonium nitrate concentration of 3 M) in order to exchange the sodium alkali cations with ammonium ions. This refluxing step is carried out four times with a fresh solution of ammonium nitrate, then the solid is filtered, washed with deionized water and dried in an oven overnight at 100°C. Finally, to obtain the zeolite in its acid form (protonated H. +) a calcination step is carried out at 550°C for ten hours (temperature rise ramp of 2°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The solid thus obtained was analyzed by X-ray diffraction and identified as consisting of IZM-2 zeolite. Characterizations using isotope NMR methods 27 AI, X-ray fluorescence and ICP provide access to the following results for IZM-2:

[0189] - weight percentage of hexacoordinated aluminum atoms AI VI : 5%, - ratio of the number of moles of silicon divided by the number of moles of aluminum in the network, in mole / mole, Si / Al: 72,

[0190] - ratio of the number of moles of sodium divided by the number of moles of network aluminum, in mole / mole, Na / AI: 0.03.

[0191] Preparation of IZM-2 / alumina support.

[0192] The IZM-2 / alumina support is obtained by mixing and extruding the IZM-2 zeolite with a Pural SB3 alumina gel. The mixed paste is extruded through a 1.8 mm diameter four-lobe die. After drying in an oven overnight at 110°C, the extrudates are calcined at 500°C for two hours (temperature rise ramp of 5°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The weight content of the IZM-2 zeolite in the support after calcination is 13% by weight.

[0193] Impregnation of platinum on IZM-2 / alumina support.

[0194] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NC>3)2. 50 grams of support are used and dry impregnated in a drageoir. After impregnation, the solid is left to mature for at least five hours in laboratory air and then left to dry overnight in an oven at 110°C. Finally, a calcination step is carried out under dry air flow (2 normal liters per hour and per gram of solid) in a tubular furnace under the following conditions:

[0195] - increase in temperature from ambient to 450°C at 5°C / min;

[0196] - two-hour stage at 450°C;

[0197] - descent to ambient.

[0198] The Pt content measured by FX on the calcined C3 catalyst is 0.29% by weight, its dispersion measured by H2 / O2 titration is 52%, its distribution coefficient measured by Castaing microprobe is 0.91.

[0199] Example 4: Preparation of a hydroconversion catalyst in accordance with the invention (C4).

[0200] Zeolite ZSM-12.

[0201] The ZSM-12 zeolite is supplied by Zeolyst. The solid was analyzed by X-ray diffraction and identified as consisting of ZSM-12 zeolite. Characterizations using isotope NMR methods 27 AI, X-ray fluorescence and ICP provide the following results for ZSM-12: - weight percentage of hexacoordinated aluminum atoms AI VI : 0% ;

[0202] - ratio of the number of moles of silicon divided by the number of moles of lattice aluminum, in mole / mole, Si / Al: 43;

[0203] - ratio of the number of moles of sodium divided by the number of moles of lattice aluminum, in mole / mole, Na / AI: 0.009.

[0204] Preparation of ZSM-12 / alumina support.

[0205] The ZSM-12 / alumina support is obtained by mixing and extruding the ZSM-12 zeolite with a Pural SB3 alumina gel. The mixed paste is extruded through a 1.8 mm diameter four-lobe die. After drying in an oven overnight at 110°C, the extrudates are calcined at 500°C for two hours (temperature rise ramp of 5°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The weight content of the ZSM-12 zeolite in the support after calcination is 6% by weight.

[0206] Impregnation of platinum on ZSM-12 / alumina support.

[0207] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NC>3)2. 50 grams of support are used and dry impregnated in a drageoir. After impregnation, the solid is left to mature for at least five hours in laboratory air and then left to dry overnight in an oven at 110°C. Finally, a calcination step is carried out under dry air flow (2 normal liters per hour and per gram of solid) in a tubular furnace under the following conditions:

[0208] - increase in temperature from ambient to 450°C at 5°C / min;

[0209] - two-hour stage at 450°C;

[0210] - descent to ambient.

[0211] The Pt content measured by FX on the calcined C4 catalyst is 0.31% by weight, its dispersion measured by H2 / O2 titration is 38%, its distribution coefficient measured by Castaing microprobe is 0.87.

[0212] Example 5: hydrotreatment using the catalyst (C1) of a feedstock from a renewable source according to a process in accordance with the invention

[0213] The hydrotreatment catalyst C1 was evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, quality of effluent separation) of the industrial process according to the invention. The different stages and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are described below.

[0214] In a reactor regulated in temperature so as to ensure isothermal operation and with a fixed bed loaded with 190 ml of hydrotreatment catalyst C1, the catalyst being previously sulfurized, the hydrotreatment of pre-refined rapeseed oil with a density of 920 kg / m is carried out. 3having an oxygen content of 11% by weight. The cetane index is 35 and the fatty acid distribution of rapeseed oil is detailed in Table 1. Prior to the hydrotreatment stage, said feedstock is added with dimethyl disulfide in order to adjust its sulfur content to 50 ppm by weight. [Table 1]

[0215] Table 1: Characteristics of the renewable rapeseed oil feedstock used as feedstock for the hydrotreatment stage.

[0216] Before hydrotreatment of the feedstock, the catalyst is sulfided in situ in the unit, with a distillation gas oil with an additive of 2% by weight of dimethyl disulfide, under a total pressure of 5.1 MPa, a hydrogen / additive gas oil ratio of 700 Nm 3 by m 3 The volume of diesel fuel added per volume of catalyst and per hour is set at 1 h' 1Sulfurization is carried out for 12 hours at 350°C, with a temperature increase ramp of 10°C per hour.

[0217] After sulfurization, the operating conditions of the unit are adjusted in order to carry out the hydrotreatment of the load:

[0218] - WH (charge volume / catalyst volume / hour): 1 h' 1 ,

[0219] - total working pressure: 5.1 MPa,

[0220] - hydrogen / charge ratio: 700 Nm 3 hydrogen / m 3 dump,

[0221] - temperature: 310°C.

[0222] The hydrogen used is supplied by Air Product and has a purity greater than 99.999% by volume.

[0223] Separation step of the effluent from the hydrotreatment step according to example 5

[0224] All of the hydrotreated effluent from the hydrotreatment step according to Example 5 is separated using a gas / liquid separator so as to recover a light fraction containing mainly hydrogen, propane, water in the form of vapor, carbon oxides (CO and CO2) and ammonia and a liquid hydrocarbon effluent consisting mainly of linear hydrocarbons. The water present in the liquid hydrocarbon effluent is removed by decantation. The liquid hydrocarbon effluent thus obtained contains an atomic oxygen content of less than 80 ppm by weight, said atomic oxygen content being measured by the infrared adsorption technique described in patent application US2009 / 0018374, and a sulfur content of 2 ppm by weight and a nitrogen content of less than 1 ppm by weight, said nitrogen and sulfur contents being measured respectively by chemiluminescence and UV fluorescence.The said liquid hydrocarbon effluent has a density of 791 kg / m. 3 The liquid hydrocarbon effluent is composed of paraffins; its composition, measured by gas chromatography, is provided in Table 2. [Table 2]

[0225] Table 2: Composition of the liquid hydrocarbon effluent used as feed for hydroconversion.

[0226] Example 6 in accordance with the invention: 2-stage hydroconversion of the liquid hydrocarbon effluent from example 5 according to a compliant process, and using catalyst (C2) in the first hydroconversion stage and catalyst (C3) in the second hydroconversion stage.

[0227] The hydroconversion catalysts C2 and C3 were evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, quality of effluent separation) of the industrial process according to the invention. The different stages and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are described below. First hydroconversion stage with catalyst (C2)

[0228] In a first reactor regulated in temperature so as to ensure isothermal operation and with a fixed bed loaded with 50 ml of hydroconversion catalyst C2, the catalyst being previously activated by sulfurization, the first hydroconversion of the liquid hydrocarbon effluent from Example 5 is carried out. Taking into account the nature of catalyst C2 (sulfide form), sulfur in the form of dialkylpentasulfide is previously introduced into said liquid hydrocarbon effluent, so as to obtain a total sulfur content of 25 ppm by weight.

[0229] Catalyst C2 undergoes an in-situ sulfurization step in the unit, with isane supplemented with 3% by weight of dimethyl disulfide, under a total pressure of 5.1 MPa, a hydrogen / isane supplemented ratio of 700 Nm 3 by m 3The volume of isane added per volume of catalyst and per hour is set at 1. Sulfurization is carried out for 12 hours at 350°C, with a temperature increase ramp of 10°C per hour.

[0230] After sulfurization, the operating conditions of the unit are adjusted in order to carry out the first stage of hydroconversion of the liquid hydrocarbon effluent in the following range of operating conditions:

[0231] - WH (charge volume / catalyst volume / hour) = 0.5 h' 1 ,

[0232] - total working pressure: 5.1 MPa,

[0233] - hydrogen / charge ratio: 350 Nm 3 hydrogen / m 3 dump.

[0234] The hydrogen used and entering the hydroconversion stage is supplied by Air Product, it has a purity greater than 99.999% by volume, it is free of hydrogen sulfide.

[0235] At the outlet of the first hydroconversion reactor, the reaction effluent is sent to a gas-liquid separation step carried out using a first flash drum operated at a pressure comparable to that of the first hydroconversion reactor. The hydrogen-rich gas phase is sent to the gas outlet of the unit. The liquid hydrocarbon phase is pressure-expanded and sent to a first stripper so as to stabilize this first hydroconverted liquid effluent; the gas phase collected at the top of the stripper is also sent to the gas outlet of the unit.

[0236] At the unit gas outlet, an online analysis by gas chromatography and a gas meter allow the mass of light hydrocarbons produced (essentially hydrocarbons with 1 to 5 carbon atoms) and present in the hydrogen flow to be calculated. The liquid effluent from the first hydroconversion is weighed separately, topped at 120°C to remove the naphtha fraction, the liquid effluent 120°C + is then reweighed and analyzed, in particular by measuring the filterability limit temperature TLF (NF EN 116) in diesel target on the first hydroconversion stage.

[0237] Temperature steps in the range 250 to 400°C were carried out in order to adjust the severity of the first hydroconversion. The measurement (typically daily) of the filterability limit temperature (operation of the first hydroconversion stage in diesel target) of the first hydroconverted liquid effluent 120°C +allows the evolution of the catalyst performance to be monitored at each temperature level. For each temperature, the test duration is extended until, on the one hand, a stable filterability limit temperature is obtained (operation of the first hydroconversion stage in diesel target) and, on the other hand, a sufficient quantity of liquid effluent 120°C+ is produced in order to feed it as a feedstock for the second hydroconversion stage.

[0238] Once the filterability limit temperature is stable, the yield in 120°C cut + (diesel) at the terminals of the first hydroconversion stage is determined according to the following calculation:

[0239] Yield 120°C + (diesel) from stage 1 = [(liquid effluent mass 120°C + ) / (load mass)] x 100, the load corresponding here to the hydrocarbon effluent from example 5.

[0240] Liquid effluent mass 120°C + corresponds to the quantity of liquid 120°C+ accumulated over a certain period of time, typically 24 hours, and the feed mass corresponds to the quantity of feed injected into the first hydroconversion reactor during the same period of time.

[0241] The temperature adjustment was carried out in order to achieve the cold property targets of a summer diesel fuel and a winter diesel fuel (target for the first hydroconversion step). For each diesel target, the characteristics obtained on the 120°C cut + stabilized from the first hydroconversion stage according to example 6, as well as the operating conditions and yields 120°C + associated (at the terminals of step 1) are reported in summary table 3.

[0242] Second hydroconversion step with catalyst (C3)

[0243] In a second reactor, temperature-controlled to ensure isothermal operation and with a fixed bed loaded with 50 ml of C3 hydroconversion catalyst, the catalyst being previously activated by reduction, the hydroconversion of the 120°C cut is carried out. + (diesel) stabilized and topped from the first hydroconversion stage according to example 6. Given the nature of the C3 catalyst (noble metal), any injection of sulfur is to be avoided.

[0244] Catalyst C3 undergoes a reduction step under hydrogen carried out at 5.1 MPa and 400°C (2-hour stage).

[0245] After reduction, the operating conditions of the unit are adjusted in order to carry out the second stage of hydroconversion of the 120°C diesel cut. + stabilized and topped from the first hydroconversion stage in the following operating conditions range:

[0246] - WH (charge volume / catalyst volume / hour) = 0.5 h'1 ,

[0247] - total working pressure: 5.1 MPa,

[0248] - hydrogen / charge ratio: 350 Nm 3 hydrogen / m 3 dump.

[0249] The hydrogen used and entering the hydroconversion stage is supplied by Air Product, it has a purity greater than 99.999% by volume, it is free of hydrogen sulfide.

[0250] At the outlet of the second hydroconversion reactor, the reaction effluent is sent to a gas-liquid separation stage carried out using a second flash drum operated at a pressure comparable to that of the second hydroconversion reactor. The hydrogen-rich gas phase is sent to the gas outlet of the unit. The liquid hydrocarbon phase is pressure-expanded and sent to a second stripper so as to stabilize this second hydroconverted liquid effluent; the gas phase collected at the top of the stripper is also sent to the gas outlet of the unit.

[0251] At the unit gas outlet, an online analysis by gas chromatography and a gas meter allow the mass of light hydrocarbons produced (essentially hydrocarbons with 1 to 5 carbon atoms) and present in the hydrogen flow to be calculated. The liquid effluent from the second hydroconversion is weighed separately, topped at 120°C to remove the naphtha fraction, the liquid effluent 120°C + is then reweighed and analyzed, in particular by measuring the disappearance point of the crystals (ASTM D5972) in kerosene target on the second hydroconversion stage.

[0252] Temperature steps in the range 250 to 400°C were carried out in order to adjust the severity of the second hydroconversion stage. The measurement (typically daily) of the crystal disappearance point (operation of the second hydroconversion stage in kerosene target) of the second liquid effluent 120°C +hydroconverted allows the evolution of the catalyst's performance to be monitored at each temperature level. For each temperature, the test duration is extended until a stable crystal disappearance point is obtained.

[0253] Once the crystal disappearance point is stable, the yield in 120°C cut + (kerosene) at the terminals of the second hydroconversion stage is determined according to the following calculation:

[0254] Yield 120°C + (kerosene) from stage 2 = [(liquid effluent mass 120°C + ) / (charge mass)] x 100, the charge corresponding here to the 120°C+ charge produced by the first hydroconversion stage and fed into the second hydroconversion stage according to example 6.

[0255] Liquid effluent mass 120°C + corresponds to the quantity of liquid 120°C +accumulated over a certain period of time, typically 24 hours, and the feed mass corresponds to the quantity of feed injected into the second hydroconversion reactor during the same period of time.

[0256] Cutting efficiency at 120°C + (kerosene) at the terminals of the process according to the invention is determined according to the following calculation:

[0257] Yield 120°C + (kerosene) of the compliant process = [(liquid effluent mass 120°C + ) / (load mass)] x 100, the load corresponding here to the hydrocarbon effluent from example 5.

[0258] Liquid effluent mass 120°C + corresponds to the quantity of liquid 120°C + accumulated over a certain period of time, typically 24 hours, and the feed mass corresponds to the quantity of feed injected into the first hydroconversion reactor during the same period of time.

[0259] The temperature adjustment was carried out in order to achieve the cold property targets of a kerosene fuel (target for this second hydroconversion step). The characteristics obtained on the 120°C cut + stabilized from the second hydroconversion stage according to example 6 in accordance, as well as the operating conditions and the associated 120°C+ yields (at the limits of stage 2 and at the limits of the compliant process) are reported in summary table 3

[0260] The reference temperature, named "Tbase" in summary table 3, is the temperature required to reach the kerosene target on the 120°C cut +produced by the non-compliant process according to Example 8 (crystal disappearance point of -50°C), the non-compliant process comprising a single hydroconversion step. The compliant process according to Example 6 was evaluated according to two variants, named respectively “Case 1” and “Case 2”. Each case has a hydroconversion rate of the liquid hydrocarbon effluent from Example 5 which varies within the first hydroconversion step.

[0261] In “Case 1”, the conversion rate of the first hydroconversion stage (catalyst C2) is adjusted to achieve a filterability limit temperature (FLT) equal to 0°C on the 120°C cut + produced, representative of a “summer diesel” fuel target. The conversion rate of the second hydroconversion stage (catalyst C3) is adjusted to achieve the kerosene target on the 120°C cut + produced by the compliant process (crystal disappearance point of -50°C).

[0262] In “Case 2”, the conversion rate of the first hydroconversion stage (catalyst C2) is adjusted to achieve a filterability limit temperature (FLT) equal to -15°C on the 120°C cut + produced, representative of a “winter diesel” fuel target. The conversion rate of the second hydroconversion stage (catalyst C3) is adjusted to achieve the kerosene target on the 120°C+ cut produced by the compliant process (crystal disappearance point of -50°C).

[0263] The process according to example 6 makes it possible to reduce the temperature used in the first hydroconversion stage; the reduction is 20°C for “Case 1” and 12°C for “Case 2”.

[0264] The compliant process according to Example 6 using a specific C3 catalyst in the second hydroconversion step makes it possible to produce a kerosene cut at a more moderate temperature than the temperature used according to the non-compliant Example 8. The reduction is 11°C for “Case 1” and 12°C for “Case 2”.

[0265] The process according to example 6 using a specific catalyst C3 in the second hydroconversion stage makes it possible to produce a kerosene cut with a yield at 120°C + overall process more favorable than that obtained according to non-compliant example 8. The gain amounts to 17.7 points for “Case 1” and 14.3 points for “Case 2”.

[0266] The use of the process according to the invention, carried out by means of two hydroconversion stages and using a specific C3 catalyst in the second hydroconversion stage, therefore allows a significant gain in the temperatures used within said process as well as a significant gain in the overall selectivity of the process in the 120°C kerosene cut. + of interest for the said process. [Table 3]

[0267] Table 3: Operating conditions, yields and properties of the cuts obtained for the different examples of the invention Example 7 in accordance with the invention: 2-stage hydroconversion of the liquid hydrocarbon effluent from example 5 according to a compliant process, and using catalyst (C2) in the first hydroconversion stage and catalyst (C4) in the second hydroconversion stage.

[0268] The hydroconversion catalysts C2 and C4 were evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, quality of separation of the effluent) of the industrial process according to the invention. The different steps and unit operations of the pilot unit mimicking the industrial implementation of the process according to the invention are identical in all respects to those described in example 6 in accordance with the invention, the only difference lies in the use of catalyst (C4) instead of catalyst (C3).

[0269] The reference temperature, named "Tbase" in summary table 3, is the temperature required to reach the kerosene target on the 120°C cut + produced by the non-compliant process according to Example 8 (crystal disappearance point of -50°C), the non-compliant process comprising a single hydroconversion step.

[0270] The process according to Example 7 was evaluated according to two variants, named respectively “Case 1” and “Case 2”. Each case presents a hydroconversion rate of the liquid hydrocarbon effluent from Example 5 which varies within the first hydroconversion stage.

[0271] In “Case 1”, the conversion rate of the first hydroconversion stage (catalyst C2) is adjusted to achieve a filterability limit temperature (FLT) equal to 0°C on the 120°C cut + produced, representative of a “summer diesel” fuel target. The conversion rate of the second hydroconversion stage (catalyst C4) is adjusted to achieve the kerosene target on the 120°C cut + produced by the compliant process (crystal disappearance point of -50°C).

[0272] In “Case 2”, the conversion rate of the first hydroconversion stage (catalyst C2) is adjusted to achieve a filterability limit temperature (FLT) equal to -15°C on the 120°C cut + produced, representative of a “winter diesel” fuel target. The conversion rate of the second hydroconversion stage (catalyst C4) is adjusted to achieve the kerosene target on the 120°C+ cut produced by the compliant process (crystal disappearance point of -50°C). The compliant process according to example 7 makes it possible to reduce the temperature used in the first hydroconversion stage; the reduction is 20°C for “Case 1” and 12°C for “Case 2”.

[0273] The compliant process according to Example 7 using a specific C4 catalyst in the second hydroconversion step makes it possible to produce a kerosene cut at a more moderate temperature than the temperature used according to the non-compliant Example 8. The reduction is 19°C for “Case 1” and 20°C for “Case 2”.

[0274] The process according to example 7 using a specific C4 catalyst in the second hydroconversion stage makes it possible to produce a kerosene cut with a yield at 120°C + overall process performance more favorable than that obtained according to non-compliant example 8. The gain amounts to 17.5 points for “Case 1” and 14.0 points for “Case 2”.

[0275] The use of the process according to the invention, carried out by means of two hydroconversion stages and using a specific C4 catalyst in the second hydroconversion stage, therefore allows a significant gain in the temperatures used within said process as well as a significant gain in the overall selectivity of the process in the 120°C kerosene cut. + of interest in the said process.

[0276] Example 8 not in accordance with the invention: One-step hydroconversion of the liquid hydrocarbon effluent from example 5 and using catalyst (C2)

[0277] The C2 hydroconversion catalyst was evaluated in a pilot unit, representative in terms of implementation (reaction operating conditions, effluent separation quality) of the non-compliant industrial process according to the invention. The various stages and unit operations of the pilot unit mimicking the industrial implementation of the non-compliant process according to the invention are described below.

[0278] In a reactor regulated in temperature so as to ensure isothermal operation and with a fixed bed loaded with 50 ml of hydroconversion catalyst C2, the catalyst being previously activated by sulfurization, the hydroconversion of the liquid hydrocarbon effluent from Example 5 is carried out. Taking into account the nature of catalyst C2 (sulfide form), sulfur in the form of dialkylpentasulfide is previously introduced into said liquid hydrocarbon effluent, so as to obtain a total sulfur content of 25 ppm by weight.

[0279] Catalyst C2 undergoes an in-situ sulfurization step in the unit, with isane added with 3% by weight of dimethyl disulfide, under a total pressure of 5.1 MPa, a hydrogen I isane added ratio of 700 Nm 3 by m 3 The volume of isane added per volume of catalyst and per hour is set at 1. Sulfurization is carried out for 12 hours at 350°C, with a temperature increase ramp of 10°C per hour.

[0280] After sulfurization, the operating conditions of the unit are adjusted in order to carry out the hydroconversion of the liquid hydrocarbon effluent in the following range of operating conditions:

[0281] - WH (charge volume / catalyst volume / hour) = 0.5 h' 1 ,

[0282] - total working pressure: 5.1 MPa,

[0283] - hydrogen / charge ratio: 350 Nm 3 hydrogen / m 3 dump.

[0284] The hydrogen used and entering the hydroconversion stage is supplied by Air Product, it has a purity greater than 99.999% by volume, it is free of hydrogen sulfide.

[0285] At the outlet of the hydroconversion reactor, the reaction effluent is sent to a gas-liquid separation stage carried out using a flash drum operated at a pressure comparable to that of the hydroconversion reactor. The hydrogen-rich gas phase is sent to the gas outlet of the unit. The liquid hydrocarbon phase is pressure-expanded and sent to a stripper so as to stabilize the hydroconverted liquid effluent; the gas phase collected at the top of the stripper is also sent to the gas outlet of the unit.

[0286] At the unit gas outlet, an online analysis by gas chromatography and a gas meter allow the mass of light hydrocarbons produced (essentially hydrocarbons with 1 to 5 carbon atoms) and present in the hydrogen flow to be calculated. The liquid effluent obtained is weighed separately, topped at 120°C to remove the naphtha fraction, the liquid effluent 120°C + is then reweighed and analyzed, in particular by measuring the disappearance point of the crystals (operation of the process not compliant with the kerosene target).

[0287] Temperature steps in the range 250 to 400°C were carried out in order to adjust the severity of the single hydroconversion step. The measurement (typically daily) of the crystal disappearance point (operation of the non-compliant process in kerosene target) of the liquid effluent 120°C +allows the evolution of catalyst performance to be monitored at each temperature level. For each temperature, the test duration is extended until a stable crystal disappearance point is obtained (operation of the process not compliant with the kerosene target). Once the crystal disappearance point is stable, the yield in the 120°C cut + (kerosene) at the terminals of the non-compliant process is determined according to the following calculation:

[0288] Yield 120°C + (kerosene) = [(liquid effluent mass 120°C + ) / (load mass)] x 100, the load corresponding here to the hydrocarbon effluent from example 5. The liquid effluent mass 120°C + corresponds to the quantity of liquid 120°C + accumulated over a certain period of time, typically 24 hours, and the feed mass corresponds to the amount of feed injected into the hydroconversion stage during the same period of time.

[0289] The temperature adjustment was carried out in order to achieve the cold property targets of a kerosene fuel (target for the non-conforming process). The characteristics obtained on the 120°C cut + stabilized according to example 8 not compliant, as well as the operating conditions and yields 120°C + associates are reported in summary table 3.

Claims

CLAIMS 1. Process for treating a feedstock from a renewable source to produce a kerosene cut comprising at least the following steps and preferably consisting of: a) a step of hydrotreating said feedstock in the presence of a fixed-bed catalyst, said catalyst comprising a hydrogenating function and an oxide support, at a temperature of between 200 and 450°C, at a pressure of between 1 and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 2000 Nm 3 of hydrogen / m 3of charge, b) a step of separating at least a portion of the effluent from step a) into at least one light gaseous fraction, at least one hydrocarbon liquid effluent, and at least one aqueous liquid effluent, c) a first step of hydroconversion of at least a portion of the hydrocarbon liquid effluent from step b) in the presence of at least one bifunctional fixed-bed hydroconversion catalyst, said catalyst comprising a hydrogenating phase containing at least one metal from group VIII and / or at least one metal from group VIB of the periodic table, an acid 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 of between 250 and 500°C, at a pressure of between 1 and 10 MPa, at an hourly space velocity of between 0.1 and 10 h -1and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3 of charge, d) a step of separating at least a portion of the effluent from step c) which makes it possible to separate at least one gaseous fraction, and at least one liquid hydrocarbon effluent, e) a step of fractionating the liquid hydrocarbon effluent from step d) into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and a heavy cut having an initial boiling point of between 250 and 300°C, or a step of stabilizing the liquid hydrocarbon effluent from step d) allowing the separation of at least one light gaseous fraction, a naphtha hydrocarbon cut and a heavy cut having an initial boiling point of between 100 and 180°C, f) a second hydroconversion step distinct from step c) of all or part of the heavy cut having an initial boiling point of between 250 and 300°C from step e) in the case where step e) is a fractionation step or of all or part of the heavy cut having an initial boiling point of between 100 and 180°C from step e) in the case where step e) is a stabilization step, said second hydroconversion step being carried out in a fixed bed in the presence of a bifunctional hydroconversion catalyst comprising at least one noble metal, said noble metal being platinum and a support comprising a silica-alumina or of a bifunctional hydroconversion catalyst comprising at least one noble metal, said noble metal being platinum and a support comprising a zeolite chosen from zeolites with structural code MTW and IZM-2, alone or as a mixture,said hydroconversion step being carried out at a temperature between 250 and 500°C, at a pressure between 1 and 10 MPa, at an hourly space velocity between 0.1 and 10 h, -1 and in the presence of a total quantity of hydrogen mixed with the charge such that the hydrogen / charge ratio is between 70 and 1000 Nm 3 / m 3of total charge treated in said step f), to produce a hydroconverted effluent, g) a step of separating at least a portion of the hydroconverted effluent from step f) which makes it possible to separate at least one gaseous fraction, and at least one liquid hydrocarbon effluent, h) a step of fractionating the liquid hydrocarbon effluent from step g) into at least one light gaseous fraction, a naphtha hydrocarbon cut, a kerosene cut having an initial boiling point of between 100 and 180°C, and a heavy cut having an initial boiling point of between 250 and 300°C, or a step of stabilizing the liquid hydrocarbon effluent from step g) allowing the separation of at least one light gaseous fraction, a naphtha hydrocarbon cut and a kerosene cut having an initial boiling point of between 100 and 180°C, said steps d) and g) and e) and h) being able to be carried out in the same equipment or in separate equipment.

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

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

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

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

6. Process according to one of claims 1 to 5 in which a catalyst used in the first hydroconversion step c) 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 sulphide form.

7. Process according to one of claims 1 to 5 in which a catalyst used in the first hydroconversion step c) comprises at least one metal from group VIII and preferably a noble metal from group VIII chosen from platinum and palladium. Preferably, the metal from group VIII of the catalyst used in step c) is platinum, preferably active in its reduced form.

8. Process according to one of claims 1 to 7 wherein in the case where the acid support of the catalyst used in the first hydroconversion step c) is one or more zeolites, said zeolites are chosen from zeolites of structural type FAU, *BEA, ISV, IWR, IWW, MEI, UWY, taken alone or in a mixture, preferably chosen from zeolites of structural type FAU and *BEA, taken alone or in a mixture, preferably, the zeolite is chosen from zeolite Y and zeolite beta taken alone or in a mixture and very preferably, the zeolite is zeolite Y and even more preferably dealuminated zeolite USY.

9. Process according to one of claims 1 to 8 in which a hydroconversion catalyst used in first hydroconversion step c) comprises nickel and tungsten, and a silica-alumina as a support, without any other binder.

10. Process according to one of claims 1 to 9 in which a second external paraffinic feed is sent to said second hydroconversion step f).

11. Process according to one of claims 1 to 10 in which a catalyst used in the second hydroconversion step f) comprises and is preferably constituted by platinum, and a support comprising and preferably constituted by a zeolite with structural code MTW and preferably ZSM-12 and an alumina binder.

12. Process according to one of claims 1 to 10 in which a catalyst used in the second hydroconversion step f) comprises and is preferably constituted by platinum, and a support comprising and preferably constituted by an IZM-2 zeolite and an alumina binder.

13. Process according to one of claims 1 to 12 in which said second hydroconversion step f) as well as steps g) and h) can be carried out on the same site as the preceding steps a) to e) of said process or on a different site, in a dedicated hydroconversion unit not belonging to said process.

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