FR2981942A1 - Method for processing load output from renewable source e.g. free fatty acids, involves allowing separated portion of effluent to undergo fractional distillation to obtain distillate fraction of liquid at moderate value - Google Patents

Abstract

The method involves separating a portion of effluent from hydrocarbon liquid, and removing a portion of water from the hydrocarbon liquid. The portion of effluent is purified in presence of hydrogen stream, where atomic oxygen content in the hydrogen stream exceeds 500 parts per million (PPM). The separated portion of effluent is allowed to undergo fractional distillation to obtain distillate fraction of liquid at a moderate value.

Description

FIELD OF THE INVENTION In an international context marked by the rapid growth in need of fuels, in particular diesel and kerosene bases in the European community, the search for new renewable energy sources that can be integrated into the traditional scheme of refining and the production of fuels is a major issue. As such, the integration into the refining process of new products of plant origin, resulting from the conversion of lignocellulosic biomass or from the production of vegetable oils or animal fats, has in recent years experienced a very lively renewed interest due to the rising cost of fossil fuels. Similarly, traditional biofuels (mainly ethanol or methyl esters of vegetable oils) have acquired a real status of supplementing petroleum fuels in fuel pools. In addition, the processes known to date using vegetable oils or animal fats are responsible for CO2 emissions, known for these negative effects on the environment. A better use of these bio-resources, such as for example their integration in the fuel pool would therefore be a definite advantage. The strong demand for diesel fuels and kerosene, coupled with the importance of environmental concerns, reinforces the interest of using renewable sources. These fillers include for example vegetable oils, animal fats, raw or having undergone prior treatment, as well as mixtures of such fillers. These fillers contain chemical structures of the triglyceride or ester or fatty acid type, the structure and the hydrocarbon chain length of the latter being compatible with the hydrocarbons present in gas oils and kerosene. One possible route is the catalytic conversion of the renewable source feedstock into deoxygenated paraffinic fuel in the presence of hydrogen (hydrotreatment). Many metal catalysts or sulfides are known to be active for this type of reaction. These hydrotreatment processes from renewable source are already well known and are described in many patents. For example, the patents are US Pat. No. 4,992,605, US Pat. No. 5,705,722, EP 1,681,337 and EP 1,741,768. The use of transition metal sulphide solids allows the production of paraffins from ester-type molecules in two reaction ways: The hydrodeoxygenation leading to the formation of water by hydrogen consumption and the formation of hydrocarbons of carbon number (Cm) 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 the formation of hydrocarbons with one less carbon (Cm_1) than the initial fatty acid chains. The liquid effluent resulting from these hydrotreatment processes, after separation, consists essentially of n-paraffins which can be incorporated into the diesel fuel and kerosene pool and is substantially free of sulfur, nitrogen and oxygen impurities. After hydrotreatment and separation of the gases, 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. In order to improve the cold properties of this hydrotreated liquid effluent, a hydroisomerization step is necessary to transform the n-paraffins into branched paraffins having better properties when cold. The patent application EP 1 741 768 describes a vegetable oil treatment process comprising a hydrotreatment followed by a hydroisomerisation step in order to improve the cold properties of the linear paraffins obtained. The catalysts used in the hydroisomerization step are bifunctional catalysts consisting of a metal active phase comprising a Group VIII metal selected from palladium, platinum and nickel, dispersed on a molecular sieve type acid support selected from the group consisting of SAPO-11, SAPO-41, ZSM-22, ferrierite or ZSM-23, said process operating at a temperature between 200 and 500 ° C, and at a pressure of between 2 and 15 MPa.

Patent Application FR 2 950 895 teaches the use of hydroisomerization catalysts comprising at least one hydro-dehydrogenating metal selected from the group formed by metals of group VIB and group VIII of the Periodic Table, taken alone or as a mixture and a composite support formed by a zeolite chosen from zeolites Y, ZSM-48, ZBM-30, IZM-1 and COK-7, taken alone or as a mixture and silicon carbide SiC. No indication of the purity of the hydrogen employed in the hydroisomerization step is provided in these patents. US2009 / 0300971 discloses a process for preparing naphtha from a feedstock from a renewable source comprising a hydrotreating step, a gas / liquid separation step followed by a hydrocracking step. The recycle hydrogen which is used in the hydrotreatment and / or hydrocracking stage can optionally be purified so as to separate the ammonia, the carbon oxides and the hydrogen sulfide before being sent into the said hydrotreatment and hydrocracking steps. It is stated that the prior removal of contaminants present in the hydrogen stream preserves the catalytic activity and selectivity of the hydrotreatment and hydrocracking catalysts. However, no precise indication of the purity of the hydrogen employed in the hydrocracking step is mentioned.

Indeed, the hydroisomerization step may use hydrogen from different sources. Depending on the nature of the different sources, the hydrogen used in the process according to the invention may or may not contain impurities. For example, a catalytic reforming unit produces hydrogen during the dehydrogenation reactions of naphthenes in aromatics and during dehydrocyclization reactions. The hydrogen produced by a catalytic reforming unit is substantially free of CO and CO2. Hydrogen can also be produced by other methods, for example by the steam reforming of light hydrocarbons or by the partial oxidation of various hydrocarbons such as heavy residues. The steam reforming process consists of converting a light hydrocarbon feed into synthesis gas, that is to say into a mixture of hydrogen (H2), carbon monoxide (CO) and carbon dioxide (CO2), and of water (H2O) by reaction with steam on a nickel-based catalyst. In this case the production of hydrogen is also accompanied by the formation of carbon oxides which are substantially eliminated by the conversion of carbon monoxide (CO) to carbon dioxide (CO2) vapor and then by elimination of CO2 by absorption for example by a solution of amines. There may also be removal of residual carbon monoxide (CO) by a methanation step. Other sources of hydrogen may also be employed such as hydrogen from catalytic cracking gases which contains significant amounts of CO and CO2. Thus, depending on its origin (s), the hydrogen used in the process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis may contain several hundred ppm by volume of carbon oxides . In attempting to develop a process for treating feedstocks from a renewable source comprising at least one hydrotreatment step and a hydroisomerization step, the Applicant has discovered that the presence of carbon monoxide (CO), of carbon dioxide (CO2 ), and more generally that the presence of molecules containing at least one oxygen atom in hydrogen, even at low levels of atomic oxygen, has a negative impact on the performance of the hydroisomerization catalyst when the paraffinic charge to be hydroisomerized is substantially free of sulfur, nitrogen and especially oxygenated impurities, such as carbon monoxide and dioxide (CO and O 2), water or alcohols and / or carboxylic acids, esters and ketones, as is the case for example after a hydrotreating step. Indeed, in the case of a paraffinic charge to be hydroisomerized which contains a significant amount of oxygenated compounds, that is to say which has not been hydrotreated, the impact of said oxygenated compounds contained in said charge is much greater than the the impact of the presence of impurities containing at least one oxygen atom in hydrogen, which then becomes negligible. Conversely, when the paraffinic feedstock is substantially free of oxygenates, the presence in hydrogen of impurities containing at least one oxygen atom has a negative impact on the hydroisomerization catalyst. An advantage of the process according to the present invention is thus to provide a process for treating a feedstock from a renewable source which undergoes a hydrotreatment before being sent to a hydrosiomerisation step which operates in the presence of a feed stream. hydrogen containing a limited content of molecules containing at least one oxygen atom, said process making it possible to improve the performance of the hydroisomerization catalyst and thus allow the improvement of the cold properties of the middle distillate cut and in particular of the the diesel fraction produced by said process, while maintaining a high cetane number. Object of the invention More specifically, the present invention relates to a process for treating a feedstock from a renewable source comprising at least: a) a step of hydrotreating said feedstock in the presence of a fixed bed catalyst, said catalyst comprising a hydro-dehydrogenating function and a supple at a temperature of between 200 and 450 ° C., at a pressure of between 1 MPa and 10 MPa, at a space velocity of between 0.1 hr -1 and 10 hr -1 and in the presence of a quantity total of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm3 of hydrogen / m 3 of feed, b) a step of separating at least a portion of the effluent from the feedstock. step a) in at least one light fraction and at least one hydrocarbon liquid effluent, c) a step of removing at least a portion of the water from the hydrocarbon liquid effluent resulting from step b), d) a step of hydroisomerization of at least a portion of the liquid hydrocarbon effluent from step c), in the presence of a fixed bed hydroisomerization catalyst, said hydroisomerization step being carried out at a temperature between 150 and 500 ° C, at a pressure of between 1 MPa and 10 MPa, at a space velocity comp between 0.1 and 10 h -1 and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm 3 / m 3 of feed, and in the presence of a hydrogen stream having undergone a purification step in the case where the atomic oxygen content in said hydrogen stream is greater than 500 ppm by volume, e) a fractionation step of the effluent from step d) for obtain at least one middle distillate fraction.

DETAILED DESCRIPTION OF THE INVENTION The present invention is particularly dedicated to the preparation of gas oil and / or kerosene fuel bases corresponding to the new environmental standards, from charges derived from renewable sources.

The feedstocks from renewable sources used in the process according to the present invention are advantageously chosen from oils and fats of vegetable or animal origin, or mixtures of such fillers, containing triglycerides and / or free fatty acids and / or esters. Vegetable oils can advantageously be crude or refined, wholly or in part, and derived from the following plants: rapeseed, sunflower, soybean, palm, palm kernel, olive, coconut, jatropha, this list not being limiting. Algae or fish oils are also relevant. Animal fats are advantageously chosen from lard or fats composed of residues from the food industry or from the catering industries.

These fillers essentially contain chemical structures of the triglyceride type which the skilled person also knows under the name tri ester of fatty acids as well as free fatty acids. A fatty acid ester tri is thus composed of three chains of fatty acids. These fatty acid chains in the form of tri-ester or in the form of free fatty acid have a number of unsaturations per chain, also called number of carbon-carbon double bonds per chain, generally between 0 and 3 but which can be higher especially for oils derived from algae which generally have a number of unsaturations in chains of 5 to 6. The molecules present in the feeds from renewable sources used in the present invention therefore have a number of unsaturations, expressed by triglyceride molecule, advantageously between 0 and 18. In these fillers, the unsaturation level, expressed as number of unsaturations per hydrocarbonaceous fatty chain, is advantageously between 0 and 6. Charges from renewable sources generally include also different impurities and especially heteroatoms such as nitrogen. Nitrogen levels in vegetable oils are generally between about 1 ppm and about 100 ppm by weight, depending on their nature.

Advantageously, the feedstock may undergo before 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 heat and / or chemical treatments well known to those skilled in the art. According to step a) of the process according to the invention, the charge, possibly pretreated, is brought into contact with a catalyst in a fixed bed at a temperature of between 200 and 450 ° C., preferably between 220 and 350 ° C. preferably between 220 and 320 ° C, and even more preferably between 220 and 310 ° C. The pressure is between 1 MPa and 10 MPa, preferably between 1 MPa and 6 MPa and even more preferably between 1 MPa and 4 MPa. The hourly space velocity, ie the volume of charge per volume of catalyst and per hour, is between 0.1 h -1 and 10 h -1. The filler is contacted with the catalyst in the presence of hydrogen. The total quantity of hydrogen mixed with the feedstock is such that the hydrogen / feedstock ratio is between 70 and 1000 Nm3 of hydrogen / m.sup.3 of feedstock and preferably between 150 and 750 Nm.sup.3 of hydrogen / m.sup.3 of feedstock. 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 Group VIII and / or Group VIB metal, taken alone or in mixture and a support selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also advantageously contain other compounds and for example oxides selected from the group formed by boron oxide, zirconia, titanium oxide, phosphoric anhydride. The preferred support is an alumina support and most preferably alumina or y.

Said catalyst is advantageously a catalyst comprising metals of group VIII preferably chosen from nickel and cobalt, taken alone or as a mixture, preferably in combination with at least one metal of group VIB, preferably selected from molybdenum and tungsten, taken alone or mixed.

The content of metal oxides of groups VIII and preferably of nickel oxide is advantageously between 0.5 and 10% by weight of nickel oxide (NiO) and preferably between 1 and 5% by weight of oxide of nickel. nickel and the content of metal oxides of groups VIB and preferably molybdenum trioxide is advantageously between 1 and 30% by weight of molybdenum oxide (MoO 3), preferably from 5 to 25% by weight, the percentages being expressed in% by weight relative to the total mass of the catalyst. The total content of metal oxides of groups VIB and VIII in the catalyst used in step a) is advantageously between 5 and 40% by weight and preferably between 6 and 30% by weight relative to the total mass. catalyst. 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 as much as possible the selectivity of the reaction towards a hydrogenation preserving the number of carbon atoms of the fatty chains. that is to say the hydrodeoxygenation route, this in order to maximize the yield of hydrocarbons entering the field of distillation of kerosenes and / or gas oils. This is why, preferably, one operates 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. Said catalyst used in step a) of hydrotreating of the process according to the invention may also advantageously contain a doping element chosen from phosphorus and boron, taken alone or as a mixture. Said doping element may be introduced into the matrix or preferably deposited on the support. It is also possible to deposit silicon on the support, alone or with phosphorus and / or boron and / or fluorine. The oxide weight content of said doping element is advantageously less than 20% and preferably less than 10% and is advantageously at least 0.001%.

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

The metals of the catalysts used in step a) of hydrotreatment of the process according to the invention are sulphide metals or metal phases and preferably sulphurized metals. It would not be departing from the scope of the present invention by using 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 downflow of liquid. Said hydrotreating step a) allows hydrodeoxygenation, hydrodenitrogenation and hydrodesulfurization of said feedstock.

According to step b) of the process according to the invention, a step of separation of at least a part and preferably all of the effluent from step a) is carried out. Said step b) makes it possible to separate at least one light fraction, at least one hydrocarbon liquid effluent. Said light fraction comprises at least one gaseous fraction which comprises the unconverted hydrogen and the gases containing at least one oxygen atom resulting from the decomposition of the oxygenated compounds during the hydrotreatment step a) and the C4- compounds, that is to say 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 liquid, and in particular to recover the hydrogen-rich gases that may also contain gases such as CO and CO2 and at least one liquid hydrocarbon effluent. Said hydrocarbon liquid effluent preferably has a sulfur content of less than 10 ppm by weight, a nitrogen content of less than 2 ppm by weight. The separation step b) can advantageously be carried out by any method known to those skilled in the art such as, for example, the combination of one or more high and / or low pressure separators, and / or distillation and / or high and / or low pressure stripping.

According to step c) of the process according to the invention, at least a portion and preferably all of the hydrocarbon liquid effluent from the separation step b) undergoes a phase of elimination of ai minus a part and preferably all of the water formed by the hydrodeoxygenation (HDO) reactions that take place during the hydrotreating step b). The purpose of this water removal step is to separate the water from the hydrocarbon liquid effluent containing the paraffinic hydrocarbons. Step c) of removing at least a portion of the water and preferably all of the water may advantageously be carried out by all methods and techniques known to those skilled in the art. Preferably, said step c) is carried out by drying, by passing on a desiccant, by flash, by decantation or by a combination of at least two of these techniques.

The atomic oxygen content of the hydrocarbon liquid effluent containing the paraffinic hydrocarbons from step c) of the process according to the invention, expressed in parts per million by weight (ppm), is preferably less than 500 ppm, preferably less than 300 ppm, very preferably less than 100 ppm by weight. The content in ppm atomic oxygen weight in said hydrocarbon liquid effluent is measured by the infrared absorption technique such as for example the technique described in US2009 / 0018374A1. According to step d) of the process according to the invention, at least part, and preferably all, of the liquid hydrocarbon effluent from step c) of the process according to the invention is hydroisomerized in the presence of a catalyst. fixed-bed hydroisomerization process, said hydroisomerization step being carried out at a temperature between 150 and 500 ° C, at a pressure of between 1 MPa and 10 MPa, at a space velocity of between 0.1 and 10 h. 1 and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm 3 / m 3 of feedstock, and in the presence of a flow of hydrogen having undergone a step of purification in the case where the atomic oxygen content in said hydrogen stream is greater than 500 ppm by volume. Preferably, said hydrogen stream undergoes a purification step in the case where the atomic oxygen content in said hydrogen stream is greater than 250 ppm by volume. Preferably, said hydrogen stream undergoes a purification step in the case where the atomic oxygen content in said hydrogen stream is greater than 50 ppm by volume. Said hydrogen flow used in the process according to the invention and preferably in step d) of the process according to the invention is advantageously generated by the processes known to those skilled in the art such as, for example, a catalytic reforming process or catalytic cracking of gases. Depending on the nature of the different sources, the hydrogen used in the process according to the invention may or may not contain impurities. The atomic oxygen content in said hydrogen stream can be measured by any method known to those skilled in the art such as, for example, by gas chromatography.

Preferably, said stream of hydrogen can be fresh hydrogen or a mixture of fresh hydrogen and recycle hydrogen, that is to say unreacted hydrogen in step d) of hydroisomerization and recycled in said step c). In the case where said stream of hydrogen contains an atomic oxygen content greater than 500 ppm by volume, preferably greater than 250 ppm by volume and preferably greater than 50 ppm by volume, said hydrogen stream undergoes a purification step prior to to be introduced in said step d). Said purification stage of the hydrogen stream can advantageously be carried out according to any method known to those skilled in the art. Preferably, said purification step is advantageously carried out according to the methods of pressure swing adsorption or PSA "Pressure Swing Adsorption" according to the English terminology, or adsorption temperature modulated or TSA "Temperature Swing Adsorption" according to the Anglo-Saxon terminology, amine washing, methanation, preferential oxidation, membrane processes, used alone or in combination. When the process uses a hydrogen recycle, a purge of the recycle hydrogen can also advantageously be carried out in order to limit the accumulation of molecules containing at least one oxygen atom such as carbon monoxide CO or carbon dioxide CO2 and thus limit the atomic oxygen content in said hydrogen stream.

According to the invention, the atomic oxygen content in said hydrogen stream used in the process according to the invention and preferably in step d) of the process according to the invention, expressed in part per million by volume (ppmv ), must be less than 500 ppmv, preferably less than 250 ppmv and very preferably less than 50 ppmv. The atomic oxygen content in said hydrogen stream is calculated from the concentrations of molecules having at least one oxygen atom in said hydrogen stream, weighted by the number of oxygen atoms present in said oxygenated molecule. By way of example, considering a stream of hydrogen containing CO from CO2, the content of atomic oxygen contained in said hydrogen stream is then equal to: ppmv (0) = ppmv (CO) + 2 * ppmv (CO2 ) with: ppmv (0) atomic oxygen content in the hydrogen stream part per million by volume, ppmv (CO) carbon monoxide content in the hydrogen stream part per million by volume, ppmv (CO2) carbon dioxide content in the hydrogen stream in parts per million by volume.

In the case where said stream of hydrogen contains an atomic oxygen content of less than 500 ppmv, preferably less than 250 ppmv and preferably less than 50 ppmv, no purification step of said hydrogen stream is carried out before said stream is introduced in said step d).

The operating conditions of the hydroisomerization step are adjusted to achieve a hydroisomerization of the non-converting charge. This means that the hydroisomerization is carried out with a conversion of the fraction 150 ° C + fraction 150 ° C- less than 20% by weight, preferably less than 10% by weight and very preferably less than 5% by weight. weight.

Preferably, step d) of hydroisomerization of the process according to the invention operates at a temperature of between 150 ° C. and 450 ° C., and very preferably between 200 ° and 450 ° C., at a pressure of between 2 ° C. and 450 ° C. MPa and 10 MPa and very preferably, between 1 MPa and 9 MPa, at an hourly space velocity advantageously 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 / hydrocarbon volume ratio is advantageously between 100 and 1000 normal cubic meters of hydrogen per cubic meter of filler and preferably between 150 and 1000 cubic meters of hydrogen per cubic meter of filler. The hydroisomerization catalyst comprises at least one hydro-dehydrogenating metal selected from the group consisting of Group VIB metals and Group VIII of the Periodic Table and at least one Bronsted acid solid, and optionally a binder. Preferably, the hydrocracking catalyst comprises either at least one noble metal of group VIII chosen from platinum and palladium, taken alone or as a mixture, active in their reduced form, or at least one non-noble metal of group VIII chosen from nickel and cobalt in combination with at least one Group VI metal selected from molybdenum and tungsten, alone or as a mixture, and preferably used in their sulfurized form. In the case where the hydrocracking catalyst comprises at least one noble metal of group VIII, the noble metal content of said hydrocracking catalyst is advantageously between 0.01 and 5% by weight relative to the finished catalyst, preferably between 0.05 and 4% by weight and very preferably between 0.10 and 2% by weight. In the case where the hydrocracking catalyst comprises at least one group VI metal in combination with at least one group VIII non-noble metal, the group VI metal content of said third hydrocracking catalyst is advantageously comprised in oxide equivalent. between 5 and 40% by weight relative to the finished catalyst, preferably between 10 and 35% by weight and very preferably between 15 and 30% by weight and the metal content of group VIII of said catalyst is advantageously included in equivalents oxide 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 metal function is advantageously introduced on the catalyst by any method known to those skilled in the art, such as for example comalaxing, dry impregnation or exchange impregnation. Preferably, the Bronsted acid solid is constituted by silica alumina or zeolite Y, SAPO-11, SAPO-41, ZSM-22, ferrierite, ZSM-23, ZSM-48, ZBM -30.1 'IZM-1, the COK-7. Optionally, a binder can advantageously also be used during the step of forming the support. A binder is preferably used when the zeolite is employed. Said binder is advantageously chosen from silica (SiO 2), alumina (Al 2 O 3), clays, titanium oxide (TiO 2), boron oxide (B 2 O 3) and zirconia (ZrO 2), 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. A preferred hydrocracking catalyst used in the process according to the invention advantageously comprises at least one noble metal, said noble metal being platinum and a silica-type alumina Bronsted acid solid, without any binder. The silica content of the silica-alumina, expressed as a percentage by weight, is generally between 1 and 95%, advantageously between 5 and 95%, and preferably between 10 and 80% and even more preferably between 20 and 70%. and between 22 and 45%. This silica content is perfectly measured using X-ray fluorescence. Several preferred catalysts used in step d) of hydrocracking of the process according to the invention are described below. A preferred catalyst used in the process according to the invention comprises a particular silica-alumina. More specifically, said catalyst comprises (and preferably consists essentially of) from 0.05 to 10% by weight and preferably from 0.1 to 5% by weight of at least one Group VIII noble metal, preferably chosen from platinum and palladium and, preferably, said noble metal being platinum, deposited on a silica-alumina support, without any binder, containing an amount of silica (SiO 2) of between 1 and 95%, expressed as a weight percentage, of 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 m2 / g, preferably between 200 m2 / g and 450 m2 / g and very preferably between 250 m2 / g and 450 m2 / g, - an average mesopore diameter of between 3 and 12 nm, preferably between 3 nm and 11 nm and very preferably e between 4 nm and 10.5 nm, - a pore volume of pores whose diameter is between the average diameter as defined above decreased by 3 nm and the average diameter as defined above increased by 3 nm is greater than 40% total pore volume, preferably between 50% and 90% of the total pore volume and very preferably between 50% and 70% of the total pore volume, - a total pore volume of between 0.4 and 1.2 ml / g, preferably between 0.5 and 1.0 ml / g and very preferably between 0.5 and 0.9 ml / g, a volume of the macropores, whose diameter is greater than 50 nm, and preferably between 100 nm and 1000 nm, representing between 5 and 60% of the total pore volume, preferably between 10 and 50% of the total pore volume and even more preferably between 10 and 40% of the total pore volume, content of alkaline or alkaline earth compounds of less than 300 ppm by weight and preferably less than at 200 ppm weight. The mean diameter of the mesopores is defined as the diameter corresponding to the cancellation of the curve derived from the mercury intrusion volume obtained from the mercury porosity curve for pore diameters between 2 and 50 m. The average diameter of the mesopores of the catalyst is advantageously measured from the porous distribution profile obtained using a mercury porosimeter. Preferably, the dispersion of the metal of said preferred catalyst is advantageously between 20% and 100%, preferably between 30% and 100% and very preferably between 40 and 100%. The dispersion, representing the fraction of metal available to the reagent relative to the total amount of metal of the catalyst, is advantageously measured, for example, by H2 / 02 titration or by transmission electron microscopy. Preferably, the distribution coefficient of the noble metal of said preferred catalyst is greater than 0.1, preferably greater than 0.2 and very preferably greater than 0.4. The distribution of the noble metal 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 the platinum poorly distributed (for example detected in a ring 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 crown, will be accessible to the reagents. The distribution coefficient of the noble metal can be measured by a Castaing microprobe. The noble metal salt is advantageously introduced by one of the usual methods used to deposit the metal on the surface of a solid. One of the preferred methods is dry impregnation which consists in introducing the metal salt into a volume of solution which is equal to the pore volume of the mass of solid to be impregnated. Before the reduction operation, the catalyst can advantageously undergo calcination, for example a treatment in dry air at a temperature of 300 to 750 ° C. and preferably at a temperature of 520 ° C., for 0.25 to 10 hours. and preferably for 2 hours. Another preferred catalyst in the process according to the invention comprises a particular second silica-alumina. More specifically, said catalyst comprises at least one hydro-dehydrogenating element chosen from the group formed by the elements of group VIB and group VIII of the periodic table, from 0.01 to 5.5% weight of oxide of an element. dopant chosen from phosphorus, boron and silicon and a non-zeolitic support based on silica-alumina containing an amount greater than 5% by weight and less than or equal to 95% by weight of silica (SiO 2), said catalyst having the following characteristics: a mean mesoporous diameter, measured by mercury porosimetry, of between 2 and 14 nm; a total pore volume, measured by mercury porosimetry, of between 0.1 ml / g and 0.5 ml / g; total porous content, measured by nitrogen porosimetry, of between 0.1 ml / g and 0.5 ml / g, a BET specific surface area of between 100 and 550 m 2 / g, and a porous volume, measured by mercury porosimetry, included in pores larger than 14 nm in diameter less than 0.1 ml / g, a porous volume, measured by mercury porosimetry, included in pores with a diameter greater than 16 nm of less than 0.1 ml / g, a porous volume, measured by mercury porosimetry, in pores with diameters greater than 20 nm, less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, contained in pores with diameters greater than 50 nm of less than 0.1 ml / g, an X-ray diffraction diagram which contains at least the principal lines characteristic of at least one of the transition aluminas included in the group composed of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina; packed filling greater than 0.75 g / ml. Another preferred catalyst used in the process according to the invention comprises (and preferably consists essentially of) 0.05 to 10% by weight and preferably 0.1 and 5% by weight of at least one noble metal of group VIII , preferably chosen from platinum and palladium and, preferably, said noble metal being platinum, deposited on a silica-alumina support, without any binder, containing an amount of silica (SiO 2) of between 1 and 95%, expressed in percentage by weight, 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 specific surface area BET of 200 to 600 m 2 / g and preferably between 250 m 2 / g and 500 m 2 / g, a mean mesopore diameter of between 3 and 12 nm, preferably between 3 nm and 11 nm and very preferably between 4 nm and 10.5 nm, - a v porous olume pores whose diameter is between the average diameter as defined above decreased by 3 nm and the average diameter as defined above increased by 3 nm is greater than 60% of the total pore volume, preferably greater than 70% of the total pore volume and very preferably greater than 80% of the total pore volume, - a total pore volume of less than 1 ml / g, preferably between 0.1 and 0.9 ml / g and very preferably between 0.2 and 0.7 ml / g, a content of alkaline or alkaline earth compounds of less than 300 ppm by weight and preferably less than 200 ppm by weight. Preferably, the dispersion of the noble metal of said preferred catalyst used in step d) of the process according to the invention is advantageously between 20% and 100%, preferably between 30% and 100% and very preferably between 40% and 100%. and 100. Preferably, the distribution coefficient of the noble metal of said preferred catalyst used in step d) of the process according to the invention is greater than 0.1, preferably greater than 0.2 and very preferably greater than 0.2. 0.4. This distribution coefficient is measured by Castaing microprobe.

Another preferred catalyst used in the process according to the invention comprises a silica-alumina and at least one Group VIIIB metal and at least one Group VIB metal, said catalyst being sulphurized. The content of these elements is advantageously comprised, in oxide equivalent, of between 5 and 40% by weight relative to the finished catalyst, preferably between 10 and 35% by weight and very preferably between 15 and 30% by weight and the Group VIII metal content of said catalyst is advantageously comprised, in oxide equivalent, of 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% by weight. and 6% by weight. The hydro / dehydrogenating metal function can advantageously be introduced on said catalyst by any method known to those skilled in the art, such as, for example, comalaxing, dry impregnation, exchange impregnation. Said hydroisomerization catalyst comprises at least one amorphous mineral support as a hydroisomerizing function, said an amorphous mineral support being chosen from silica-aluminas and silica aluminas and preferably silica aluminas, containing an amount greater than 5% by weight and less than or equal to 95% by weight of silica (SiO 2).

According to step e) of the process according to the invention, the hydroisomerized effluent from step d) undergoes a fractionation step, preferably in a distillation train which incorporates an atmospheric distillation and optionally a vacuum distillation, to obtain at least one middle distillate fraction.

Said step e) is intended to separate the gases from the liquid, and in particular to recover the hydrogen-rich gases that may also contain light such as the C1-C4 cut and at least one gas oil cut, at least one kerosene cut and at least one less a naphtha cut. The recovery of the naphtha fraction is not the subject of the present invention, but this section can advantageously be sent to a steam cracking or catalytic reforming unit.

The products, diesel base and / or kerosene, obtained according to the process according to the invention are endowed with excellent characteristics. The diesel base obtained is of excellent quality: its sulfur content is less than 10 ppm by weight, its total aromatics content is less than 5% by weight, and the polyaromatic content is less than 2% by weight. cetane is excellent, greater than 55, the density is less than 840 kg / m 3, and most often less than 820 kg / m 3, its kinematic viscosity at 40 ° C is 2 to 8 mm 2 / s, its properties The cold-holding properties are compatible with the standards in force, with a filterability limit of less than -15 ° C and a cloud point below -5 ° C. The kerosene obtained has the following characteristics: a density of between 775 and 840 kg / m 3, a viscosity at -20 ° C. of less than 8 mm 2 / s, a point of disappearance of crystals of less than -47 ° C. a flash point above 38 ° C, a smoke point greater than 25 mm.

EXAMPLES Example 1 Preparation of a Hydroisomerization Catalyst According to the Invention (C2) The silica-alumina powder is prepared according to the synthesis protocol described in patent FR 2,639,256 (Example 3). The amounts of orthosilicic acid and aluminum alkoxide are chosen so as to have a composition of 70% by weight of Al 2 O 3 alumina and 30% by weight of SiO 2 silica in the final solid. The dried powder is brought into contact with an aqueous solution of nitric acid, the amount of nitric acid being 5% by weight relative to the amount of powder and the amount of aqueous solution such as loss on ignition at 550 ° C. C of the cake obtained is about 60% by weight. This cake is kneaded and extruded. The kneading is done on a Z-arm kneader. The extrusion is carried out by passing the dough through a die provided with orifices 1.4 mm in diameter. The extrudates thus obtained are dried in an oven at 110 ° C. and then calcined under a dry air flow rate (ramp up to 5 ° C./min). The calcination temperature is adjusted to obtain a specific surface area of 299 m 2 / g.

The silica-alumina extrudates are then subjected to a dry impregnation step with an aqueous hexachloroplatinic acid solution H2PtC16, left to mature in a water-based maturator for 24 hours at room temperature and then calcined for two hours under dry air in a bed. crossed at 500 ° C (ramp temperature rise of 5 ° C / min). The weight content of platinum of the finished catalyst after calcination is 0.47%.

The characteristics of the catalyst thus prepared are the following: a mean diameter of the mesopores of 6.5 nm, a pore volume of the pores whose diameter is between the average diameter as defined previously decreased by 3 nm and the average diameter such as as defined previously increased by 3 nm equal to 60% of the total pore volume, - a total pore volume of 0.70 ml / g, - a volume of macropores, whose diameter is greater than 50 nm representing 29% of the total pore volume a BET surface area of 299 m 2 / g, a sodium content of 102 13 ppm by weight, a noble metal dispersion of 85% and a noble metal distribution coefficient of 0.95. EXAMPLE 2 Hydroconversion of an N-Hexadecane Model Molecule on a Hydroisomerization Catalyst According to the Invention (C2) The Impact of the Presence of CO and CO2 in Hydrogen Was First Evaluated in Hydroconversion of Hydrogenation a model paraffin, n-hexadecane. The conversion of n-hexadecane on the catalyst C2 is studied on a micro-unit of catalytic test working in crossed bed. Typically 1 gram of catalyst C 2 milled and then sieved with the fraction 250 - 315 microns is charged into the reactor. The catalyst is then dried at atmospheric pressure under a stream of nitrogen for one hour at 150 ° C. (6NiN.sub.2 / gram of catalyst / h) and then reduced at atmospheric pressure under a stream of hydrogen for 2.5 hours at 450.degree. (6 NI H2 / gram of catalyst / h). The operating conditions are then adjusted to 335 ° C, a total pressure of 50 bar and n-hexadecane is injected into the unit. The operating conditions are the following: total pressure: 50 bar, temperature: 335 ° C., hourly space velocity of n-hexadecane: 7 h -1, hydrogen / n-hexadecane ratio: 1080 N / I, use different standard mixtures can vary the content of CO or CO2 in hydrogen. All chemicals used are commercial, n-hexadecane from Sigma-Aldrich (anhydrous, purity 99-F), hydrogen and nitrogen from Air Liquide (grade U, 02 <10 ppmv, H2O <40 ppmv ) and the different standard CO / H2 and CO2 / H2 mixtures from Air Liquide as well. After passing through the reactor, the effluent is then depressurized, cooled and the liquid and gaseous effluents are analyzed using a Varian 3400 GC type gas chromatograph equipped with a flame ionization detector (FID). ) and a Supelco Petrocol DH 50.2 capillary column 50.2 m long, 0.2 mm internal diameter and 0.5 μm film thickness. Good separation is achieved by column temperature programming of 45-220 ° C (2 ° C-1), for liquid effluent analysis, and between 45 ° C and 75 ° C (2 ° C). .min-1) for the analysis of the gas phase. The pressure of hydrogen (carrier gas) at the top of the column is 11.5 psi (input splitting rate of the injection of 40 cm.sup.-1). The injector and detector temperatures are 300 ° C. Quantification of the hydrocarbons in the gaseous effluent is performed using ethane as an external standard. The yield of the isomerization products of n-hexadecane is calculated as follows: Yield, (% wt) = (z x 100 LiAcn sum of the surfaces for all the hydrocarbons with n carbon atoms (liquid effluent) 16 X merged liAis0_c16 sum of the area of the corresponding peaks Ac to the iso-hexadecanes measured mass hydrocarbon (liquid effluent) mgaz mass hydrocarbons (effluent gas) mpesée + M gases The analyzes carried out make it possible to follow the stabilization of the catalyst C2 over time and only the The results obtained after stabilization of the catalyst are considered.The stabilization time of the catalyst is extremely fast, of the order of a few hours.

Table 1 below thus reports the evolution of the isomerization yield obtained on the catalyst C2 for different levels of oxygenated compounds in hydrogen. The results illustrate the negative impact of oxygenated molecules present in hydrogen on the production of isomerization products of n-hexadecane on the catalyst C2. The limitation of the content of oxygenated compounds that may be present in the hydrogen used during the hydroisomerization step is therefore of obvious interest. Table 1: evolution of the isomerization yield during the hydroconversion of n-hexadecane on the catalyst C2 for different contents of oxygenated compounds in hydrogen. Oxygen content in H2O <40 431 ppm 431 ppm 1077 ppm 1077 ppm hydrogen / ppmv 02 <10 CO CO2 CO CO2 oxygen content <60 431 862 1077 2154 atomic in hydrogen / ppmv Yield iso / wt% 40 , 3 34.7 27.7 29.9 23.1 Example 3: Preparation of a hydrotreating catalyst (Cl) The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with levels of MoO3 molybdenum of 16% by weight, nickel oxide NiO 3% by weight and phosphorus oxide P2O5 of 6% by weight relative to the total weight of the finished catalyst, supplied by AXENS.

EXAMPLE 4 Hydrotreatment and Hydroisomerization of a Filler Resulting from a Renewable Source According to a Process not According to the Invention Step a): Hydrotreatment In a Temperature-Controlled Reactor in a Way to Ensure Isothermal and Fixed-Bed Operation Loaded with 190 ml of Hydroprocessing catalyst C1, the catalyst being previously sulphurized, is introduced 170 g / h pre-refined rapeseed oil density 920 kg / m3 with an oxygen content of 11% by weight and sulfur of 15 ppm by weight. The cetane number is 35 and the fatty acid distribution of the rapeseed oil is detailed in Table 2. Table 2: Characteristics of rapeseed oil used as feed for hydrotreatment Fatty acid composition (% ) 14: 0 0.1 16: 0 5.0 16: 1 0.3 17: 0 0.1 17: 1 0.1 18: 0 1.5 18: 1 trans <0.1 18: 1 cis 60 , 1 18: 2 trans <0.1 18: 2 cis 20.4 18: 3 trans <0.1 18: 3 cis 9.6 20: 0 0.5 20: 1 1.2 22: 0 0.3 22: 1 0.2 24: 0 0.1 24: 1 0.2 700 Nm3 of hydrogen / m3 of feed are introduced into the reactor maintained at a temperature of 300 ° C, at a hourly space velocity of 1 hr -1 and at a pressure of 5 MPa.15 Step b): separation of the effluent from step a) All of the hydrotreated effluent from step a) is separated using a gas separator / liquid so as to recover a light fraction containing predominantly hydrogen, propane, water in vapor form, carbon oxides (CO and CO2) and ammonia and a hydrocarbon effluent liquid boné 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 11 ppm weight and a nitrogen content of less than 1 ppm by weight. Said liquid hydrocarbon effluent then undergoes the hydroisomerization stage c. Step c): Hydroisomerization of the liquid hydrocarbon effluent from step b) The hydroisomerization step is also carried out in fixed bed passed from the hydrocarbon base from step c). The operating conditions of this hydroisomerization step are described below: WH (charge volume / volume of catalyst / hour) = 1 h -1 total working pressure: 5 MPa - hydrogen ratio / charge: 700 normal liters / liter The temperature is adjusted so as to have a conversion of the fraction 150 ° C + fraction 150 ° C-less than 5% by weight during hydroisomerization. Before testing, the catalyst undergoes a reduction step under the following operating conditions: hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst - ambient temperature rise 120 ° C: 10 ° C / min one-hour stage at 120 ° C. rise from 120 ° C. to 450 ° C. at 5 ° C./min two-hour stage at 450 ° C. pressure: 0.1 MPa The flow of hydrogen used in the hydroisomerisation stage is derived from a standard mixture of Air Liquide now containing 800 ppmv of CO, or 800 ppmv of oxygen, said hydrogen stream not undergoing any purification step. The effluent from step c) is then fractionated by distillation to determine the average distillate yield (150 ° C + cut). Chromatographic analysis also gives access to the yield of branched hydrocarbons, indicated in the "% isomerization" table, in the middle distillate cut. Moreover, the filterability limit temperature and the cetane number of the middle distillate cut are determined. The main characteristics of the effluent produced are reported in Table 3.

Table 3: Main characteristics of the effluents produced by hydroisomerisation Yield (% weight) 150 ° C cut - 3 150 ° C + cut (Diesel) 97% isomerization in the middle distillate cut 74 Fuel properties (150 ° C + cut) Cetane number ( ASTMD613) 87 Filterability limit temperature (° C) -2 Example 5: Hydrotreating and hydroisomerization of a feedstock from a renewable source according to a process according to the invention The hydrocarbon liquid effluent from the separation step b) used for the hydroisomerization step according to Example 5 is identical to that of Example 4. The operating conditions of this hydroisomerization step are identical to those described in Example 4. Nevertheless, the hydrogen used in this step has a much lower oxygen content: The flow of hydrogen sent in the hydroisomerization step is a flow of hydrogen of Air Liquide U quality, that is to say comprising t in 02 <10 ppmv, and H20 <40 ppmv. The atomic oxygen content in said hydrogen stream is therefore less than 60 ppmv. The effluent from the hydroisomerization step is then fractionated by distillation to determine the average distillate yield (150 ° C + cut). Chromatographic analysis also gives access to the yield of branched hydrocarbons. Moreover, the filterability limit temperature and the cetane number of the middle distillate cut are determined. Table 4: Main characteristics of effluents produced by hydroisomerisation Yield (% wt) 150 ° C - 4 cut 150 ° C + cut (Diesel) 96% isomerization in the middle distillate cut 80 Fuel properties (150 ° C + cut) Cetane number ( ASTMD613) Filterability Limit Temperature (° C) -16 In view of the results reported in Tables 3 and 4, the limitation of the oxygen content in hydrogen induces an improvement in the performance of the hydroisomerization catalyst. This has the consequence of increasing the yield of branched distillate cut-off hydrocarbons, which implies a marked improvement in the filterability limit temperature and thus in the cold properties, while maintaining a high cetane number.

Claims (8)

REVENDICATIONS1. A process for treating a feedstock from a renewable source comprising at least: a) a step of hydrotreating said feedstock in the presence of a fixed bed catalyst, said catalyst comprising a hydro-dehydrogenating function and an amorphous support, at a temperature between 200 and 450 ° C, at a pressure of between 1 MPa and 10 MPa, at a space velocity of between 0.1 h-1 and 10 h-1 and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm3 of hydrogen / m 3 of feed, b) a step of separation of at least a portion of the effluent from step a) in at least one light fraction and at least one hydrocarbon liquid effluent, c) a step of removing at least a portion of the water from the liquid hydrocarbon effluent from step b), d) a step of hydroisomerization of at least a portion of the hydrocarbon liquid effluent resulting from the step c), in the presence of a fixed bed hydroisomerization catalyst, said hydroisomerization step being carried out at a temperature between 150 and 500 ° C, at a pressure of between 1 MPa and 10 MPa, at a space velocity between 0.1 and 10 h -1 and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 70 and 1000 Nm 3 / m 3 of feed, and in the presence of a hydrogen stream having undergone a purification step in the case where the atomic oxygen content in said hydrogen stream is greater than 500 ppm by volume, e) a fractionation step of the effluent from step d) for obtain at least one middle distillate fraction. 2. Process according to claim 1, in which the filler derived from renewable sources is chosen from oils and fats of vegetable or animal origin, or mixtures of such fillers, containing triglycerides and / or free fatty acids and / or esters. 3. Method according to one of claims 1 or 2 wherein the feedstock is contacted with a fixed bed catalyst at a temperature between 220 and 350 ° C, at a pressure is between 1 MPa and 6 MPa, at a space velocity of between 0.1 h-1 and 10 h-1. The feed is brought into contact with the catalyst in the presence of hydrogen and in the presence of a total amount of hydrogen mixed with the feed such that the hydrogen / feed ratio is between 150 and 750 Nm 3 of hydrogen / m3 feedstock. 4. Method according to one of claims 1 to 3 wherein the separation step b) is carried out by the combination of one or more high and / or low pressure separators, and / or distillation steps and / or or high and / or low pressure stripping. 5. Method according to one of claims 1 to 4 wherein said step c) is carried out by drying, by passing on a desiccant, flash, decantation or a combination of at least two of these techniques. 6. Method according to one of claims 1 to 5 wherein said flow of hydrogen undergoes a purification step in the case where the atomic oxygen content in said hydrogen stream is greater than 250 ppm by volume. 7. Method according to one of claims 1 to 6 wherein said flow of hydrogen undergoes a purification step in the case where the atomic oxygen content in said hydrogen stream is greater than 50 ppm by volume. 8. Method according to one of claims 1 to 7 wherein said purification step is carried out according to pressure swing adsorption methods or PSA "Pressure Swing Adsorption" according to the English terminology, or adsorption temperature-modulated or TSA "Temperature Swing Adsorption" according to the English terminology, amine scrubbing, methanation, preferential oxidation, membrane processes, used alone or in combination.20