Improved process for the production of kerosene from synthetic hydrocarbons, comprising steps arranged in series and having low shift conversion

Abstract

The aim of the present invention is to provide a process for the production of kerosene from a paraffin feedstock produced by Fischer-Tropsch synthesis, wherein the process comprises at least one hydroconversion step and at least one hydroisomerisation step arranged in series, wherein the operating conditions and the catalysts used in the steps are specifically selected to selectively produce a kerosene fraction, while minimising the production of light fractions such as C1-C4 light gases and naphtha, and the production of oil, wherein the latter is limited to a possible purge in an effluent recycle.

Description

[0001] Description

[0002] IMPROVED PROCESS FOR PRODUCING KEROSENE FROM SYNTHETIC HYDROCARBONS IN SERIES WITH LOW CONVERSION PER PASS

[0003] Field of invention

[0004] The Fischer-Tropsch process produces synthetic hydrocarbons from a gaseous feedstock composed mainly of hydrogen and carbon monoxide, also called synthesis gas (CO+H2).

[0005] In particular, in the low-temperature Fischer-Tropsch process, the synthesis gas is catalytically transformed into water, oxygenates, olefins and essentially linear paraffins also called n-paraffins, in gaseous, liquid or solid form under standard conditions. After separation of the water, the resulting synthetic hydrocarbons are mainly composed of predominantly linear paraffins but also contain olefins and oxygenates. The effluent from the Fischer-Tropsch process, at the outlet of the Fischer-Tropsch synthesis unit, can be divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes.

[0006] These synthetic hydrocarbons produced cannot be directly incorporated into conventional fuel pools or used as lubricants. For example, the melting point of a normal paraffin with 14 carbon atoms and a boiling point of 254°C is approximately 6°C, which makes its direct incorporation into the kerosene fraction impossible given the required specification of -40°C for the crystal disappearance point for kerosene prepared from paraffins produced by Fischer-Tropsch synthesis according to ASTM D7566. A hydroisomerization step is necessary to lower the pour points of the various hydrocarbon compounds. Furthermore, processing steps are necessary to remove olefinic and oxygenated compounds, increase the yield of middle distillates and improve the properties of the produced fractions in order to meet their specifications.To carry out the hydroisomerization of Fischer-Tropsch synthesis products, so-called bifunctional catalysts are generally used. They are characterized by the association of an acid function and a hydrogenating function. For the heavy fraction, that is to say the fraction having an initial boiling point above the maximum temperature imposed by the specifications for the use of kerosene, the cracking activity of the aforementioned catalysts will be sought in order to convert the molecules into molecules with shorter chains and therefore with a lower boiling point, until the products formed integrate the desired distillation interval. For the light fraction, that is to say covering a distillation interval adapted to the targeted one, it will only be a question of using the isomerizing activity of the aforementioned catalysts.To do this, we can then choose to play on the formulation of bifunctional catalysts in order to make them more or less selective towards cracking or towards isomerization, but this may prove insufficient with a propensity of so-called isomerizing catalysts to overcrack the products to the point of forming compounds that are too light (gas or gasoline) which are then no longer usable depending on the desired objective, i.e. kerosene.

[0007] In order to illustrate the problem, we will cite below some results from the prior art allowing it to be evaluated.

[0008] First of all, patent application FR3084082 describes a process for producing middle distillates, i.e. diesel and kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis, characterized by the treatment of two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes. In this process, fractionation of the heavy wax fraction makes it possible to form two other more or less heavy fractions, the lighter of the two then being mixed with the cold condensate fraction before this mixture is hydrotreated and isomerized. In parallel with this treatment, the heavy fraction called waxes is hydrocracked to form an effluent which is finally fractionated together with the effluent of the hydroisomerized fraction to obtain at least one gasoline fraction and at least one middle distillate fraction.Although the scheme uses a specific hydroisomerizing catalyst based on at least one IZM-2 zeolite, it appears that approximately 22% by weight of the Fischer-Tropsch synthesis effluent is in the form of naphtha and it can also be assumed that by seeking to specifically produce kerosene without diesel, the quantity of naphtha would be even higher, which does not allow the problem of maximizing the kerosene yield to be addressed at the desired specifications.

[0009] Patent application WO15063213 discloses a process for converting a paraffinic feedstock, which may be a feedstock from Fischer-Tropsch synthesis, in which the feedstock is transformed in a step a) to obtain a feedstock that is at least partially isomerized on a sequence of two catalysts in series, a first catalyst 1 and a second catalyst 2 in which catalyst 2 is more isomerizing and less crunchy than catalyst 1, then a step of separating the isomerized effluent into one or more middle distillate fractions and a residual fraction.

[0010] The description specifies that the feedstock transformation step can be carried out in one or more reactors. The description then describes a fractionation step which can lead to a diesel fraction, a kerosene fraction and a residual fraction boiling above the middle distillates. Catalysts 1 and 2 are preferably composed of an amorphous acid support and preferably of the platinum / silica-alumina type for the first and a molecular sieve and preferably of the platinum / ZSM-12 zeolite type with a silicic binder for the second.

[0011] An optional vacuum distillation step of the residual fraction can be implemented to obtain a base oil distillate fraction and a second residual fraction boiling between 450 and 550°C. A portion of said second residual fraction can be recycled into the hydrocracking and hydroisomerization reaction step. The base oil distillate fraction can be sent to a catalytic dewaxing step. It will be noted that in the case where step a) is implemented in two reactors in series, the entire feedstock from the first reactor is sent to the second reactor. The use of the silica binder for the preparation of the more isomerizing catalyst 2 is presented as essential in achieving selectivities for diesel and / or kerosene.However, the examples show that despite the partial conversion (about 20% by weight of base oil is not recycled), 10% by weight of gas is formed and only 67 to 71% of middle distillates are produced, the distinction between diesel and kerosene not being indicated. It therefore appears to us that the proposed solution does not address the problem posed of maximizing the kerosene yield to the desired specifications.

[0012] Patent application WO05001006 also describes a process for treating a hydrocarbon wax, which may originate from the Fischer-Tropsch synthesis, during which this feedstock is successively subjected to hydrocracking and then to hydroisomerization before fractionation to separate the middle distillate from the heavy cut, the latter also being separated into a base oil having a high viscosity index. The examples described and based on a preferred catalytic system based on a first nickel-tungsten sulfide catalyst supported on silica-alumina, then a platinum noble metal catalyst supported on SAPO-11 reveal fairly low middle distillate yields with the objective of maximizing base oil production. Here again, the implementation does not make it possible to address the problem of maximizing the kerosene yield to the desired specifications.

[0013] Patent application WO14001546 again covers a process for producing diesel, heavy distillate and residual base oil from a feedstock derived from Fischer-Tropsch synthesis comprising a hydroconversion / hydroisomerization step to obtain an at least partially isomerized feedstock followed by a step of separating the isomerized feedstock into a diesel, a heavy distillate and a residual fraction having a T10 of between 200 and 540°C. In the disclosed process, a portion of the residual fraction is recycled to the hydroconversion / hydroisomerization step, the other portion undergoing a dewaxing step to obtain a base oil. The examples show that with such a process, the production of distillates, in this case diesel, remains low, a significant proportion of the feedstock not being converted into a lighter cut and being reserved for the production of base oil, almost 50% by weight.

[0014] This is also the case of patent application WO005003067, which describes a process for producing base oil from a feedstock produced by Fischer-Tropsch synthesis, which first comprises separating the feedstock into a cut boiling in the middle distillate range and below, a heavy cut, and an intermediate cut. Subsequently, the intermediate fraction is treated in a hydroisomerization step to produce base oils, preferably over a catalyst that may be based on ZSM-12. In parallel, the heavy fraction is sent to a hydroconversion step to produce lighter compounds, while the fraction not converted at this step is treated in a hydroisomerization step to produce base oils.Here again, the examples highlight the significant proportion of heavy fraction, base oil, produced by the process, approximately 20% by weight, but also the light part, naphtha or gas, also approximately 20% by weight, which does not allow us to respond to the problem posed of maximizing the kerosene yield to the desired specifications.

[0015] Finally, patent application WO14001552 also describes a process for transforming the effluents from a Fischer-Tropsch synthesis to produce three fractions, the first two being middle distillates and the third being a base oil. To do this, the aforementioned synthesis effluent is separated into a light fraction and a heavy fraction. The heavy fraction is then subjected to hydrocracking to form a distillate, a heavy distillate and a heavy residual fraction. Subsequently, the distillate and heavy distillate fractions undergo dewaxing, but so does the heavy residual fraction, the latter forming a base oil, which does not address the problem addressed by maximizing the kerosene yield to the desired specifications.

[0016] Application WO09041478 discloses a process for producing diesel fuel comprising firstly a step of fractionating the effluent from Fischer-Tropsch synthesis into two fractions, an intermediate fraction and a heavier wax fraction, followed, on the one hand, by a step of hydroisomerization of the intermediate fraction, and on the other hand, by a step of hydrocracking the wax fraction. Finally, the process described comprises another fractionation step applied to the mixture of the hydroisomerized intermediate fraction and the heavy fraction of hydrocracked wax into two fractions including a kerosene fraction and a diesel fraction, the latter then being combined in adjusted proportions to obtain target properties. This process makes it possible to guarantee the achievement of precise specifications, but it does not maximize the yield since a fraction of kerosene or diesel will necessarily have to be set aside.

[0017] The research work carried out by the applicant led him to discover that in order to selectively produce a kerosene cut with properties that comply with the specifications in force, in particular for the cutting points and cold properties such as the freezing point, it is necessary to judiciously combine at least two hydroconversion and hydroisomerization stages arranged in series while ensuring that operating conditions are maintained such that in these two stages the hydroconversion and hydroisomerization take place with minimal selectivity towards the gas and naphtha cuts.

[0018] An objective of the present invention is therefore to provide a process for producing kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis comprising at least one hydroconversion step and at least one hydroisomerization step arranged in series, the operating conditions and the catalysts used in said steps being chosen in a very specific manner, to selectively produce a kerosene cut, while minimizing the production of light cuts such as C1-C4 light gases, naphtha and the production of oil, the latter being limited to a possible purge on an effluent recycle.

[0019] Subject of the invention

[0020] The present invention relates to a process for producing kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis comprising at least the following steps and preferably consists of: a) hydrotreatment of at least a portion and preferably all of the paraffinic feedstock produced by Fischer-Tropsch synthesis in the presence of a hydrotreatment catalyst and operating at a temperature of between 250 and 450°C, at a pressure of between 0.5 and 15 MPa, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 3000 normal liters of hydrogen per liter of feedstock, and at an hourly volumetric velocity of between 0.1 and 40 h 1, b) fractionating in one or more stages the effluent from stage a) into at least one gaseous fraction, a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 100 and 160°C, at least one kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C, and a heavy fraction having an initial boiling point above 250°C, c) hydrocracking all of said heavy fraction having an initial boiling point above 250°C from stage b) to produce a hydrocracked effluent, in the presence of a hydrocracking catalyst and operating at a temperature of between 250 and 450°C, at a pressure of between 0.2 and 15 MPa, at a rate spatial between 0.1 and 10 h -1, and at a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed, said hydrocracking catalyst comprising at least one noble metal from group VIII of the periodic table and a support comprising at least one acidic solid, said operating conditions of said step c) being adjusted so that the conversion per pass into products having boiling points greater than or equal to a temperature Tx into products with boiling points lower than this same temperature Tx is less than 0.5, preferably less than 0.4 and preferably less than 0.3 and more preferably less than 0.2, Tx being between 250 and 320°C, preferably between 270 and 310°C and very preferably between 280 and 300°C, d) fractionating in one or more stages at least all or part of the effluent from step c), into at least one fraction gaseous,a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 100 and 160°C, at least one kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C, and a residual heavy fraction having an initial boiling point above 250°C, said fractionation steps b) and d) being advantageously able to be carried out in the same fractionation unit or in separate fractionation units,e) the hydroisomerization of at least a portion of the kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C resulting from fractionation step b) and of all or part of the kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C resulting from fractionation step d) in the case where said fractionation steps b) and d) are distinct, in the presence of a hydroisomerization catalyst operating at a temperature of between 200 and 450°C, a pressure of between 1 and 15 MPa, a space velocity of between 0.1 and 10 h, 1, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed, to produce a hydroisomerized effluent, said hydroisomerization catalyst comprising at least one noble metal from group VIII of the periodic table and a support comprising at least one binder and at least one zeolite chosen from zeolites of structural type TON, FER, EUO, AEL, *MRE, MTW, MOR, *BEA and zeolite IZM-2, taken alone or as a mixture, f) recycling all or part of the residual heavy fraction having an initial boiling point above 250°C in hydrocracking step c), g) optionally purging part of the hydrocracked effluent from hydrocracking step c), h) fractionating all or part of the hydroisomerized effluent from step e) hydroisomerization, into at least one gaseous fraction,a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 120 and 160°C, and at least one kerosene fraction having an initial boiling point of between 120 and 160°C.,

[0021] The process according to the invention also applies to a situation in which the Fischer-Tropsch synthesis effluent has previously been divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes, and where these two fractions must first be combined before their simultaneous hydrotreatment.

[0022] Generally speaking, the said process applies to all or part of the effluent from Fischer Tropsch synthesis.

[0023] An advantage of the present invention is therefore to provide a process for producing kerosene from a paraffinic effluent produced by Fischer-Tropsch synthesis which makes it possible both to maximize the production of kerosene and to limit the production of light cracked products which cannot be incorporated into a kerosene pool. Another advantage of the present invention is to provide a process for producing kerosene which makes it possible to improve operational flexibility with respect to variations in the feedstock in terms of the proportions between condensates and waxes in said feedstock.

[0024] In the remainder of the text, and unless otherwise specified, the term "heavy fraction" corresponds to a hydrocarbon effluent obtained at the end of the hydrotreatment and fractionation step(s) after said hydrotreatment step, said hydrocarbon effluent having an initial boiling point greater than 250°C, advantageously between 250 and 320°C, preferably between 270 and 310°C and more preferably between 280 and 300°C.

[0025] In the remainder of the text, and unless otherwise specified, the term "residual heavy fraction" corresponds to a hydrocarbon effluent obtained at the end of the hydrocracking and fractionation step(s) after said hydrocracking step, said hydrocarbon effluent having an initial boiling point greater than 250°C, advantageously between 250 and 320°C, preferably between 270 and 310°C and more preferably between 280 and 300°C.

[0026] The recycling of at least part of the residual heavy fraction to the hydrocracking stage is essential to the invention because this makes it possible to crack all or part of said residual heavy fraction and to upgrade it in particular into a kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C. Thus, the non-recycled residual heavy fraction is reduced to the strict minimum, preferably it is zero and, failing that, it corresponds to a purge.

[0027] In the following text, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81 ème edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification, and group VIB to the metals of column 6.

[0028] 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.

[0029] In the present description, the expression "greater than..." is understood as strictly greater than, and symbolized by the sign ">", and the expression "less than" as strictly less than, and symbolized by the sign "<". n detailed description of the invention

[0030] According to the invention, the present invention relates to a process for producing kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis.

[0031] The Fischer-Tropsch process, well known to those skilled in the art, makes it possible to produce synthetic hydrocarbons from a gaseous feedstock composed mainly of hydrogen and carbon monoxide, also called synthesis gas (CO+H2).

[0032] The synthesis gas feeding the Fischer-Tropsch process can advantageously be produced from natural gas, coal, biomass, any source of hydrocarbon compounds or a mixture of these sources. It can also come from any existing process for supplying CO2, such as capture processes applied to industrial fumes or biogenic CO2 and to which a step of transformation of CO2 into CO has been associated. In addition, hydrogen can also come from any existing process known to those skilled in the art. These include reforming, but also water electrolysis processes. Hydrogen can also be produced by other methods such as steam reforming of light hydrocarbons or by the partial oxidation of different hydrocarbons such as heavy residues.Other sources of hydrogen can also be used, such as hydrogen from catalytic cracking gases, which contain significant amounts of CO and CO2. The hydrogen used can also come from the outlet gas of a hydrotreatment unit; in this case, this hydrogen can undergo more or less advanced purification steps to remove impurities such as ammonia (NH3) or hydrogen sulfide (H2S).

[0033] In particular, in the low-temperature Fischer-Tropsch process, the synthesis gas is catalytically transformed into water, oxygenates, olefins and essentially linear paraffins also called n-paraffins, in gaseous, liquid or solid form under standard conditions. After separation of the water, the resulting synthetic hydrocarbons are mainly composed of paraffins but also contain olefins and oxygenates.

[0034] Preferably, after separation of the water, the paraffinic effluent from the Fischer-Tropsch process comprises an n-paraffin content greater than 70% by weight and even more preferably greater than 80% by weight relative to the total mass of said effluent.

[0035] Preferably, the paraffinic effluent from the Fischer-Tropsch process is free from heteroatomic impurities other than oxygen, such as, for example, sulfur, nitrogen or metals.

[0036] Generally, the paraffinic effluent from the Fischer-Tropsch process is, at the outlet of the Fischer-Tropsch synthesis unit, divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes. However, in the case of the present invention, this separation is not obligatory. The process can therefore be applied both to a so-called global effluent grouping the cold condensate and the waxes as well as to a part of each of these effluents or to a reconstituted cut of these two effluents.

[0037] The light fraction called cold condensate corresponds to hydrocarbons in the gaseous state under the conditions of the Fischer-Tropsch reaction and the heavy fraction called waxes corresponds to hydrocarbons in the liquid state under the conditions of the Fischer-Tropsch reaction.

[0038] The light fraction of cold condensates has an initial boiling point T1 of between 15 and 40°C and preferably between 20 and 35°C and a final boiling point T2 of between 350 and 400°C and preferably between 360 and 380°C and preferably less than 370°C.

[0039] The heavy wax fraction advantageously has an initial boiling point T3 of between 100 and 300°C and preferably between 125 and 200°C. Preferably, said heavy fraction contains paraffins with a number of carbon atoms greater than 8.

[0040] This heavy fraction has the particularity of being solid at room temperature due to its composition, namely essentially n-paraffins with a high number of carbon atoms of up to 150.

[0041] In the light fraction, the paraffin content is greater than 70% by weight, the olefin content less than 20% by weight and the oxygenated compound content less than 10% by weight, the contents being expressed as a percentage by weight relative to the total mass of said light fraction. In the heavy fraction, the paraffin content is greater than 80% by weight, the olefin content less than 15% by weight and the oxygenated compound content less than 5% by weight, the contents being expressed as a percentage by weight relative to the total mass of said heavy fraction.

[0042] According to the invention, the process comprises a) step of hydrotreatment of at least part and preferably all of the effluent from the Fischer-Tropsch synthesis in the presence of a hydrotreatment catalyst and operating at a temperature of between 250 and 450°C, at a pressure of between 0.5 and 15 MPa, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 3000 normal liters of hydrogen per liter of feed, and at an hourly volumetric flow rate of between 0.1 and 40 h 1 .

[0043] Preferably, said hydrotreatment step a) operates at a temperature of between 280 and 420°C and preferably between 330 and 390°C, at a pressure of between 1 and 9 MPa, a hydrogen flow rate adjusted to obtain a ratio of between 250 and 2000 normal liters of hydrogen per liter of feed and preferably between 500 and 1500 normal liters of hydrogen per liter of feed, and at an hourly volumetric flow rate of between 0.25 and 20 h 1 and preferably between 0.5 and 10 h 1 .

[0044] Said hydrotreatment step makes it possible to reduce the content of olefinic and unsaturated compounds as well as to decompose the oxygenated compounds present in the effluent resulting from the Fischer-Tropsch synthesis. The hydrotreatment catalyst used in step a) is a conventional hydrotreatment catalyst. Said catalyst comprises at least one metal from group VIII and / or group VI of the periodic table of elements. Preferably, the catalyst comprises at least one metal from the group of metals formed by nickel, molybdenum, tungsten, cobalt, ruthenium, indium, palladium and platinum alone or as a mixture, and comprises at least one support chosen from aluminas, boron oxides, magnesia, zirconia, titanium oxides and clays or a combination of these oxides, preferably said support is an alumina.Said catalysts can advantageously be prepared by any methods known to those skilled in the art or can be purchased from companies specializing in the manufacture and sale of catalysts. The shape of the supports and catalysts involved in step a) can be spherical, or extruded in the form of cylinders, trilobes or even quadrilobes.

[0045] In the case of the use of non-noble metals of group VIII, a combination of at least one metal of group VI, preferably molybdenum or tungsten and at least one metal of group VIII, preferably cobalt or nickel of the periodic table of elements is advantageously used. The concentration of non-noble metal of group VIII, when used, is advantageously from 0.01% to 15% by weight of oxide equivalent relative to the finished catalyst and that of the metal of group VI is advantageously from 5% to 40% by weight of oxide equivalent relative to the finished catalyst. When a combination of metals of group VI and group VIII is used, the catalyst is then preferably used in a reduced or sulfurized form.

[0046] Under these conditions, the content of unsaturated and oxygenated molecules in the effluent from said step a) is reduced to less than 0.5% by weight and to approximately less than 0.1% by weight in general. The hydrotreatment step is carried out under conditions such that the conversion into products having boiling points greater than or equal to 300°C into products having boiling points lower than 300°C is limited to 20% by weight, preferably is less than 10% by weight and even more preferably is less than 5% by weight.

[0047] The effluent from step a) can advantageously be sent to a step for removing at least part of the water formed during said step a) and preferably all of the water formed before being sent to step b) of fractionation according to the invention.

[0048] Said step of removing at least a portion of the water can advantageously be carried out by any methods and techniques known to those skilled in the art, for example by drying, passing through a desiccant, flash or decantation. The process according to the invention comprises a step b) of fractionation in one or more stages of the effluent from step a) into at least one gaseous fraction advantageously comprising the light gases C1-C4, a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 100 and 160°C, at least one kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C, and a heavy fraction having an initial boiling point above 250°C.

[0049] Said fractionation step can advantageously be carried out by atmospheric distillation and in certain cases by the combination of atmospheric distillation and vacuum distillation.

[0050] Another embodiment of this fractionation step consists of stripping the effluent from step a) prior to atmospheric fractionation in order to purge the gaseous fractions at the inlet of said atmospheric fractionation column.

[0051] In a preferred embodiment, two kerosene fractions are separated during fractionation step b), a light kerosene fraction with an initial boiling point of between 100 and 160°C and a final boiling point of between 160 and 220°C, and a heavy kerosene fraction with an initial boiling point of between 160 and 220°C and a final boiling point of between 250 and 320°C. In this embodiment, only the heavy kerosene fraction is treated in said hydroisomerization step e). The light kerosene fraction is advantageously not sent to said hydroisomerization step e) according to the invention.

[0052] The light kerosene fraction from fractionation step b) can then advantageously be mixed with the kerosene fraction from fractionation step h), the mixture then constituting the final kerosene cut obtained by said process and having the required specifications.

[0053] This embodiment has the advantage of reducing the size of the hydroisomerization reactor in step e) and also the size of the fractionation unit in step h).

[0054] According to the invention, the process comprises a step c) of hydrocracking at least a part and preferably all of the heavy fraction from the fractionation step b) to produce a hydrocracked effluent, in the presence of a hydrocracking catalyst and operating at a temperature of between 250 and 450°C, at a pressure of between 0.2 and 15 MPa, at a space velocity of between 0.1 and 10 h 1, and with a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed, said hydrocracking catalyst comprising at least one noble metal from group VIII and a support comprising an acidic solid, said operating conditions of said step c) being adjusted so that the conversion per pass into products having boiling points greater than or equal to a temperature Tx into products with boiling points lower than this same temperature Tx is less than 0.5, preferably less than 0.4, preferably less than or equal to 0.3 and preferably less than 0.2, Tx being between 250 and 320°C, preferably between 270 and 310°C and very preferably between 280 and 300°C.

[0055] Preferably, step c) operates at a temperature between 280 and 450°C, and even more preferably between 320 and 420°C, at a pressure between 0.5 and 10 MPa, more preferably between 1 and 9 MPa, at a space velocity between 0.2 and 7 h 1 , and more preferably between 0.5 and 5 h 1 , and with a hydrogen flow rate adjusted to obtain a ratio of between 150 and 1500 normal liters of hydrogen per liter of charge and more preferably between 300 and 1500 normal liters of hydrogen per liter of charge.

[0056] The feedstock of step c) advantageously has an initial boiling temperature greater than 250°C, preferably between 250 and 320°C, preferably between 270 and 310°C and more preferably between 280 and 300°C.

[0057] According to the invention, the hydrocracking catalyst used during step c) comprises at least one noble hydro-dehydrogenating metal from group VIII of the periodic table and a support comprising at least one acidic solid and preferably a Bronsted acidic solid, and optionally a binder.

[0058] Preferably, said hydrocracking catalyst comprises at least one noble metal from group VIII chosen from platinum and palladium, taken alone or in a mixture, preferably active in their reduced form.

[0059] The noble metal content of said catalyst is advantageously between 0.01 and 5% by weight relative to the finished catalyst, preferably between 0.02 and 4% by weight, very preferably between 0.03 and 3% by weight, even more preferably between 0.05 and 2% by weight and very preferably between 0.05 and 1% by weight.

[0060] 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.

[0061] Advantageously, the Bronsted acid solid comprises and preferably consists of silica alumina and / or zeolite.

[0062] In the case where the Bronsted acid solid comprises a zeolite, this is preferably a zeolite of structural type FAU, *BEA, ISV, IWR, IWW, MEI, UWY, taken alone or in a mixture and preferably chosen from zeolites of structural type FAU and *BEA, taken alone or in a mixture. In a preferred embodiment, the zeolite is chosen from zeolite Y and zeolite beta taken alone or in a mixture and preferably the zeolite is zeolite Y and very preferably dealuminated zeolite USY. Preferably the Bronsted acid solid comprises and is preferably made of silica-alumina.

[0063] Optionally, said hydrocracking catalyst support may also comprise a binder. Preferably, said support comprises a binder when said support comprises a zeolite. Said binder is advantageously chosen from silica (SiO2), alumina (AI2O3), clays, titanium oxide (TiO2), 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.

[0064] A preferred hydrocracking catalyst according to the invention advantageously comprises at least one noble metal, said noble metal being platinum, and a silica-alumina as a Bronsted acid solid, without any other binder.

[0065] 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.

[0066] A preferred hydrocracking catalyst used in 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 (SiO2) 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:

[0067] - 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,

[0068] - an average diameter of the mesopores measured by mercury porosimetry of between 4 and 12 nm, preferably between 4 and 11 nm and very preferably between 5 and 11 nm,

[0069] - a total pore volume measured by mercury porosimetry of between 0.2 and 1.2 ml / g, preferably between 0.3 and 1.0 ml / g and very preferably between 0.3 and 0.9 ml / g,

[0070] - a volume of macropores measured by mercury porosimetry, the diameter of which is greater than 50 nm, less than 0.02 ml / g.

[0071] - a content of alkali or alkaline-earth compounds of less than 300 ppm by weight and preferably less than 200 ppm by weight. 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.

[0072] 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.

[0073] 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.

[0074] 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 of 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 such as for example a treatment in dry air at a temperature of 300 to 750°C and preferably at a temperature equal to 450°C, for 0.25 to 10 hours and preferably for 2 hours.

[0075] During this step c) the feed entering the reactor undergoes, in contact with the catalyst and in the presence of hydrogen, essentially hydrocracking reactions which, accompanied by hydroisomerization reactions of the n-paraffins, will make it possible to improve the quality of the products formed and more particularly the cold properties of the kerosene, and also to selectively produce kerosene compared to naphtha, which requires maintaining low conversion rates per pass.

[0076] According to the invention, the operating conditions of said step c) are adjusted so that the conversion per pass into products having boiling points greater than or equal to a temperature Tx into products with boiling points lower than this same temperature Tx is less than 0.5, preferably less than 0.4, preferably less than 0.3 and more preferably less than 0.2, Tx being between 250 and 320°C, preferably between 270 and 310°C and very preferably between 280 and 300°C. The temperature Tx, and the conversion rate resulting therefrom, is (are) preferably chosen so as to maximize the kerosene yield of the process according to the invention while respecting the specifications of aviation kerosene.

[0077] The intended effect of the operating conditions applied and the conversion per pass they generate is to limit overcracking into light products such as gas or naphtha, the latter being unwanted.

[0078] When the temperature increase required to maintain the target pass conversion becomes too high and the selectivity to light products, gas and / or gasoline increases, the catalytic cycle can be interrupted and the catalyst is discharged and then replaced by a new fresh catalyst with which the expected effect will be found again.

[0079] The conversion per pass is defined as:

[0080] Tx Conversion + in Tx- = [(% mass of Tx- effluent) - (% mass of Tx- load)] / [100 - (% mass of Tx- load)], with:

[0081] - % mass of Tx- effluent: mass percentage of compounds having boiling points lower than Tx in the hydrocracking effluent,

[0082] - % mass of Tx- charge: mass percentage of compounds having boiling points lower than Tx in the hydrocracking charge.

[0083] It is advantageously found that the fraction of compounds having a boiling point in the naphtha range, therefore lower than those which can be incorporated into the kerosene product which is found in the hydrocracked effluent from step c) is then greatly reduced by limiting the conversion of the heavy fraction corresponding to the compounds with a boiling point higher than Tx to levels lower than 0.5, preferably lower than 0.4, preferably lower than or equal to 0.3 and preferably lower than 0.2.

[0084] The process according to the invention comprises a step d) of fractionating at least all or part of the effluent from step c).

[0085] Preferably, all of the effluent from step c) is fractionated in step d).

[0086] In another embodiment, the effluent from step c) undergoes a purge before being fractionated in step d).

[0087] Step d) allows the fractionation of said effluent into at least one gaseous fraction advantageously comprising light C1-C4 gases, at least one gasoline or naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 100 and 160°C, at least one kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C, and a residual heavy fraction having an initial boiling point above 250°C, said fractionation steps b) and d) being advantageously able to be carried out in the same fractionation unit or in separate units.

[0088] Said fractionation step can advantageously be carried out by atmospheric distillation and in certain cases by the combination of atmospheric distillation and vacuum distillation.

[0089] Another embodiment of this fractionation step consists of stripping the effluent from step c) prior to atmospheric fractionation in order to purge the gaseous fractions at the inlet of said atmospheric fractionation column.

[0090] Said fractionation steps b) and d) can advantageously be carried out in the same fractionation unit or in separate units. Preferably, fractionation steps b) and d) are carried out in the same unit.

[0091] In the preferred embodiment where fractionation steps b) and d) are carried out in the same fractionation unit, all or part of the effluents from step a) and step c) are treated jointly in said same fractionation unit. In this case, fractionation steps b) and d) are combined.

[0092] In a preferred embodiment in which fractionation steps b) and d) are carried out in the same unit, the effluent from step c) undergoes purging before being fractionated.

[0093] According to the invention, the process comprises a step e) of hydroisomerization of at least a portion of the kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C resulting from the fractionation step b) and of all or part of the kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C resulting from the fractionation step d) in the case where said fractionation steps b) and d) are distinct. Step e) is carried out in the presence of a hydroisomerization catalyst operating at a temperature of between 200 and 450°C, a pressure of between 1 and 15 MPa, a space velocity of between 0.1 and 10 h 1 , a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed to produce a hydroisomerized effluent.

[0094] Preferably, said hydroisomerization step e) operates at a temperature of between 250 and 450°C and preferably between 280 and 430°C, a pressure of between 1 and 10 MPa and preferably between 1 and 9 MPa, a space velocity of between 0.2 and 7 h 1 and preferably between 0.5 and 5 hours 1 , a hydrogen flow rate adjusted to obtain a ratio of between 150 and 2000 normal liters of hydrogen per liter of charge and preferably between 150 and 1500 normal liters of hydrogen per liter of charge.

[0095] According to a first variant, all of the kerosene fractions obtained after step(s) b) and d) are sent to the hydroisomerization step e).

[0096] According to a second variant, in the case where fractionation steps b) and d) are carried out in separate fractionation units, two kerosene fractions are advantageously separated during fractionation step d), a light kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of less than or equal to 250°C, preferably less than 230°C and very preferably less than 200°C and a heavy kerosene fraction having an initial boiling point greater than 200°C, and a final boiling point of between 250 and 320°C.

[0097] In this embodiment, only the heavy kerosene fraction from fractionation step d) is sent to said hydroisomerization step e), mixed with the kerosene fraction from fractionation step b). The light kerosene fraction is advantageously not sent to said hydroisomerization step e) according to the invention and is mixed directly with the kerosene fraction from final fractionation step h).

[0098] According to a second variant, in the case where fractionation steps b) and d) are carried out in a single fractionation unit, two kerosene fractions are advantageously separated during the fractionation step, a light kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of less than or equal to 250°C, preferably less than 230°C and very preferably less than 200°C and a heavy kerosene fraction having an initial boiling point of greater than 200°C, and a final boiling point of between 250 and 320°C.

[0099] In this embodiment, only the heavy kerosene fraction is sent to said hydroisomerization step e). The light kerosene fraction is advantageously not sent to said hydroisomerization step e) according to the invention and is mixed directly with the kerosene fraction from the final fractionation step h).

[0100] The catalyst used in hydroisomerization step e) according to the invention is advantageously of the bifunctional type, that is to say that it has a hydro / dehydrogenating function and a hydroisomerizing function.

[0101] According to the invention, the hydroisomerization catalyst comprises and preferably consists of at least one noble metal from group VIII of the periodic table (which acts as a hydro / dehydrogenating function), and a support comprising and preferably consisting of at least one binder and at least one zeolite chosen from zeolites of structural type TON, preferably chosen from ZSM-22 and NU-10, taken alone or as a mixture, zeolites of structural type FER preferably chosen from ZSM-35 and ferrierite, taken alone or as a mixture, zeolites of structural type EUO preferably chosen from EU-1 and ZSM-50, taken alone or as a mixture, zeolites of structural type AEL preferably SAPO-11, zeolites of structural type *MRE preferably chosen from ZSM-48, ZBM-30, EU-2 and EU-11 , taken alone or in mixture, zeolites of MTW structural type preferably chosen from ZSM-12, TPZ-12, Theta-3, NU-13,CZH-5, taken alone or in a mixture, zeolites of structural type MOR preferably chosen from mordenite and LZ-21 1, taken alone or in a mixture, zeolites of structural type *BEA preferably chosen from Beta and Tschernichite) and zeolite IZM-2, taken alone or in a mixture.,

[0102] The hydro / dehydrogenating function

[0103] Preferably the noble metal from group VIII of the catalyst used in step e) is chosen from palladium and platinum and is preferably platinum.

[0104] Advantageously, the hydro / dehydrogenating (metallic) element and preferably platinum can be introduced onto the catalyst support by any method known to those skilled in the art, such as for example co-mixing, dry impregnation, exchange impregnation.

[0105] According to one or more embodiments, the content of noble metal from group VIII, and preferably the platinum content, in the catalyst used in step e) is between 0.01% and 4% by weight, preferably between 0.05% and 2% by weight, relative to the total weight of said catalyst.

[0106] The catalyst used in step e) may also advantageously further comprise at least one additional metal chosen from the group formed by the metals of groups IIIA, IVA and VIIB 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.

[0107] Preferably, the content of at least one additional metal in the catalyst used in step e) is between 0.01% and 2% by weight, preferably between 0.05% and 1% by weight, relative to the total weight of said catalyst.

[0108] According to one embodiment, the sulfur content in the hydroisomerization catalyst is such that the ratio of the number of moles of sulfur to the number of moles of the at least one metal from group VIIIB 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 hydroisomerization 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.

[0109] The hydroisomerizing function. According to the invention, the catalyst contains at least one zeolite of structural type TON (eg chosen from ZSM-22 and NU-10, taken alone or in a mixture) and / or FER (eg chosen from ZSM-35 and ferrierite, taken alone or in a mixture) and / or EUO (eg chosen from EU-1 and ZSM-50, taken alone or in a mixture) and / or AEL (eg SAPO-11) and / or *MRE (eg chosen from ZSM-48, ZBM-30, EU-2 and EU-11, taken alone or in a mixture) and / or MTW (eg chosen from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or in a mixture) and / or MOR (eg chosen from mordenite or LZ-211, taken alone or in a mixture) and / or *BEA (e.g. chosen from Beta or Tschernichite). The structural codes are defined in the classification of the International Zeolite Association (IZA: http: / / www.iza-structure.org / databases / ). The zeolite can also be IZM-2, whose structural code is not known.

[0110] Preferably, the catalyst contains at least one zeolite chosen from IZM-2 zeolite and a zeolite of MTW structural type preferably chosen from ZSM-12, TPZ-12, Theta-3, NU-13, CZH-5, taken alone or as a mixture.

[0111] Preferably, the catalyst contains at least one zeolite chosen from IZM-2 and ZSM-12, taken alone or as a mixture.

[0112] The zeolites are preferably essentially in acid form, that is to say that the atomic ratio between the monovalent compensation cation (for example sodium) and the aluminum inserted in the crystal lattice of the solid is advantageously less than 0.1, preferably less than 0.05 and very preferably less than 0.01. According to one or more embodiments, the zeolites entering into the composition of said hydroisomerization catalyst are advantageously calcined. According to one or more embodiments, said zeolites are exchanged by at least one treatment with a solution of at least one ammonium salt so as to obtain the ammonium form of the zeolites which, once calcined, leads to the acid form of said zeolites.

[0113] According to a preferred embodiment, the catalyst comprises a support comprising an IZM-2 zeolite alone or a support comprising a ZSM-12 zeolite alone.

[0114] Zeolite IZM-2 is a microporous crystallized solid whose crystal structure and preparation process are described in patent application FR2918050A1. The structural code of zeolite IZM-2 is not known to date. Zeolite ZSM-12 is a microporous crystallized solid whose crystal structure is described on the website of the International Zeolyst Association (http: / / www.iza-structure.org / ). It is a one-dimensional zeolite with 12 MR, its structural code is MTW. A preparation process for this zeolite is for example described in the article Synthesis of zeolite ZSM-12 in the system (MTEA)2O-Na2O-SiO2-Al2O3-H2O by S. Ersnt et al. (Zeolites, 7, 5, 458-462, DO110.1016 / 0144-2449(87)90015-7).

[0115] Preferably, the catalyst used in step e) comprises a zeolite content of between 1% and 90% by weight, preferably between 3% and 80% by weight, and more preferably between 4% and 60% by weight, preferably between 4% and 30% by weight and more preferably between 4% and 20% by weight relative to the total weight of said catalyst.

[0116] The binder.

[0117] Preferably, the binder of the catalyst support of step e) is amorphous or crystallized. Preferably, the binder used in the catalyst support of step e) is advantageously chosen from the group formed by alumina, silica, silica-alumina, clays, titanium oxide, boron oxide, zirconia and aluminates, taken alone or as a mixture. Preferably, the binder is alumina. Preferably, said binder may contain alumina in all its forms known to those skilled in the art, such as for example alpha, gamma, eta, delta type aluminas.

[0118] Preferably, the catalyst used in step e) comprises a binder content of between 10% and 99% by weight, relative to the total weight of said catalyst Le., so as to ensure the addition to 100% by weight of the elements constituting the catalyst used in step e).

[0119] According to the invention, the catalyst support comprises the zeolite mixed with a binder. The shaping of the support in the form of a mixture is preferably carried out by co-mixing, extrusion and then heat treatment of the zeolite with the binder or a precursor of the binder, such as for example boehmite, which by heat treatment is transformed into alumina.

[0120] A preferred catalyst of step e) comprises and preferably consists of platinum, and a support comprising and preferably consists of a ZSM-12 zeolite and an alumina binder.

[0121] Another preferred catalyst of step e) comprises and preferably consists of platinum, and a support comprising and preferably consists of an IZM-2 zeolite and an alumina binder.

[0122] According to a preferred embodiment, the catalyst used in step e) more particularly comprises, and preferably consists of:

[0123] - 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;

[0124] - from 0.01% to 4% by weight, preferably from 0.05% to 2% by weight of at least one metal from group VIIIB, preferably platinum;

[0125] - 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 VIIB;

[0126] - optionally a sulfur content, preferably such that the ratio of the number of moles of sulfur to the number of moles of group VIIIB metal(s) is between 0.3 and 3; and

[0127] - optionally at least one binder, preferably alumina, ensuring the 100% complement in the catalyst, relative to the total weight of the catalyst of step e). Preferably, the catalyst used in step e) 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 e) is shaped in the form of crushed powders, tablets, rings, balls, wheels. Techniques other than extrusion, such as pelletizing or coating, can advantageously be used.The shaping can also advantageously be carried out in the presence of the various constituents of the catalyst and extrusion of the mineral paste obtained, by pelletizing, shaping in the form of balls with a rotating beader or drum, drop coagulation, "oil-drop", "oil-up", or any other known process of agglomeration of a powder containing alumina and possibly other ingredients chosen from those mentioned above.

[0128] Furthermore, the supports used in the method according to the present invention may advantageously have been treated, as is well known to those skilled in the art, with additives to facilitate shaping and / or improve the final mechanical properties of the supports. As examples of additives, mention may in particular be made of cellulose, carboxymethylcellulose, carboxyethylcellulose, tall oil, xanthan gums, surfactants, flocculating agents such as polyacrylamides, carbon black, starches, stearic acid, polyacrylic alcohol, polyvinyl alcohol, biopolymers, glucose, polyethylene glycols, etc.

[0129] Extrusion can advantageously be carried out by any conventional, commercially available tool. The paste resulting from the mixing is advantageously extruded through a die, for example using a piston or a single-screw or twin-screw extrusion machine. This extrusion step can advantageously be carried out by any method known to those skilled in the art.

[0130] The catalyst support used in step e) according to the present invention is then advantageously subjected to a drying step carried out according to any technique known to those skilled in the art.

[0131] Preferably, the drying is carried out under an air flow. Said drying may also advantageously be carried out under a flow of any oxidizing, reducing or inert gas. Preferably, the drying is advantageously carried out between 50 and 180°C, preferably between 60 and 150°C and very preferably between 80 and 130°C.

[0132] Said support, possibly dried, then preferably undergoes a calcination step.

[0133] Said calcination step is advantageously carried out in the presence of molecular oxygen, for example by carrying out an air sweep, at a temperature advantageously greater than 200°C and less than or equal to 1100°C. Said calcination step can advantageously be carried out in a crossed bed, a licked bed or in a static atmosphere. For example, the furnace used can be a rotating rotary kiln or a vertical kiln with radial crossed layers. Preferably, said calcination step is carried out between more than one hour at 200°C and less than one hour at 1100°C. The calcination can advantageously be carried out in the presence of water vapor and / or in the presence of an acidic or basic vapor. For example, the calcination can be carried out under partial pressure of ammonia. Post-calcination treatments can optionally be carried out, so as to improve the properties of the support, for example textural properties.

[0134] Preferably, the noble metal contained in said catalyst used in step e) can advantageously be reduced. One of the preferred methods for carrying out the reduction of the metal is treatment at a temperature between 150 and 650°C and a total pressure between 0.1 and 25 MPa. For example, a reduction may comprise a two-hour stage at 150°C followed by a temperature increase to 450°C at a rate of 1°C / min followed by a two-hour stage at 450°C; during the reduction step, the hydrogen flow rate may be 1000 normal m 3 hydrogen / m 3 catalyst and the total pressure can be kept constant at 0.1 MPa. Any reduction method can advantageously be considered, either in situ (the reduction of the catalyst is carried out in the same unit where the catalytic reaction is carried out), or ex situ (the reduction is carried out outside the unit where the catalytic reaction is carried out, before loading the catalyst into the unit).

[0135] The operating conditions implemented in step e) make it possible to obtain an isomerized effluent containing little or no olefinic compounds and oxygenated compounds.

[0136] The isomerized effluent from step e) is then sent to fractionation step h) to obtain at least one kerosene fraction.

[0137] According to the invention, the process comprises the recycling of all or part of the residual heavy fraction having an initial boiling point above 250°C in hydrocracking step c).

[0138] In the variant where steps b) and d) are carried out in a single fractionation unit, the hydrocracked effluent from step c), optionally purged, is advantageously recycled in said common fractionation unit in a mixture with all or part of the effluent from hydrotreatment step a).

[0139] In this case, said common fractionation unit makes it possible to obtain at least one heavy fraction which comprises the heavy fraction of the hydrotreated effluent and the residual heavy fraction of the hydrocracked effluent.

[0140] In said variant, said heavy fraction which is sent to the hydrocracking step c) comprises the heavy fraction of the hydrotreated effluent and the residual heavy fraction of the hydrocracked effluent.

[0141] In the variant where steps b) and d) are carried out in separate fractionation units, only the hydrocracked effluent from step c), optionally purged, is fractionated in said fractionation step d) to obtain at least one residual heavy fraction having an initial boiling point above 250°C. Preferably, all or part of said residual heavy fraction, optionally purged, is advantageously recycled directly into said hydrocracking step c).

[0142] The recycling of the residual heavy fraction is a key step of the invention because this recycling step allows at least part of the residual heavy fraction to be converted into at least one kerosene fraction. This allows the kerosene yield of the process to be increased. It is also more advantageous to operate at low conversion per pass and with a recycle than to attempt to convert more of the heavy fraction in a single step and without recycle since the latter possibility would lead to overcracking the product and thus forming a significant quantity of gas and / or gasoline.

[0143] Said recycling is advantageously carried out with a rate corresponding to 1 to 19 times the mass flow rate of fresh feed from hydrotreatment step a) and separated in step b). Preferably, the recycle rate is between 2 and 9 and even more preferably, the recycle rate is carried out between 3 and 6.

[0144] The recycle rate is defined as the ratio of the mass flow rate of feed entering hydrocracking stage c), i.e. the cumulative flow rate of the heavy fraction and the residual heavy fraction to the flow rate of the heavy fraction. It can be calculated directly from the conversion per pass previously defined according to the following formula:

[0145] Recycle Tx = 1 / (Conversion of Tx + in Tx)

[0146] The recycling rate is then advantageously between 2 and 20, preferably between 3 and 10 and even more preferably between 4 and 6.

[0147] An optional purge of a portion of the hydrocracked effluent from step c) may be carried out. The purge constitutes a fraction which will not be recycled in any step of the process according to the invention and in particular neither to the hydrocracking step c) nor to the hydroisomerization step e).

[0148] Generally, it is a part of the residual heavy fraction and advantageously the heaviest fraction of said residual heavy fraction.

[0149] The implementation of the purge makes it possible to maximize the quantity of residual heavy fraction recycled to the hydrocracking stage c) while maintaining the performance of the process, any heavy compounds refractory to hydrocracking likely to be accumulated during the recycle stages being able to be eliminated by the purge.

[0150] Said purging can advantageously be carried out upstream or downstream of the fractionation step and preferably downstream of the fractionation step d) when steps b) and d) are carried out in separate fractionation units. In another embodiment, the purging is carried out during, upstream or downstream of the fractionation step and preferably downstream of the fractionation step when steps b) and d) are carried out in the same fractionation unit.

[0151] Preferably, the purge corresponds to less than 10% by weight of the total charge entering step a) of hydrotreatment, very preferably, less than 7% by weight, very preferably, less than 5% by weight and even more preferably less than 3% by weight.

[0152] The purge can be carried out by any method known to those skilled in the art and can, for example, consist of a draw-off or a simple flash, the separation temperature then being chosen to be as high as possible, preferably greater than 300°C, preferably greater than 370°C, very preferably greater than 450°C and even more preferably greater than 540°C.

[0153] The effluent from step c), deducted purge, is then fractionated.

[0154] The process according to the invention comprises a step h) of fractionation of all or part of the effluent from the hydroisomerization step e) to obtain at least one gaseous fraction advantageously comprising light C1-C4 gases, at least one naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 120 and 160°C, and at least one kerosene fraction having an initial boiling point of between 120 and 160°C.

[0155] At the end of this fractionation step h), the kerosene obtained has all the specifications required for use as aviation fuel according to current standards.

[0156] In the preferred embodiment where fractionation steps b) and d) are carried out in the same fractionation unit, and in the case where a light kerosene fraction comes from said fractionation unit, said light kerosene fraction is mixed with the kerosene fraction coming from final fractionation step h).

[0157] In the preferred embodiment where fractionation steps b) and d) are carried out in separate fractionation units, and in the case where one or more light kerosene fractions are derived from said fractionation steps b) and / or d), said light kerosene fraction(s) is / are mixed with the kerosene fraction resulting from the final fractionation step h).

[0158] The process according to the invention makes it possible to obtain a final kerosene fraction, consisting of the kerosene fraction resulting from fractionation step h) only or optionally consisting of said kerosene fraction resulting from fractionation step h) in a mixture with a light kerosene fraction according to the preferred embodiments described above.

[0159] A mass balance allows to demonstrate an excellent selectivity of the process towards this final kerosene cut, the unwanted cuts being produced in reduced proportions compared to the other implementations described in the literature to our knowledge. In particular, the yield of the cuts having a boiling point higher than the target end point of the kerosene was reduced to the minimum as required by the invention, the purge being less than 10% by weight of the mass of feed entering step a). Furthermore, the yield of the light cuts resulting from the cracking during one or other of steps a), c) or e) is also extremely reduced due to the serial treatment of the heavy fraction of the incoming feed first on a hydrocracking catalyst with a low conversion rate per pass then on a hydroisomerization catalyst selective towards isomerization at the expense of cracking.

[0160] Description of figures

[0161] [FIG .1] Figure 1 illustrates embodiment no. 1 in which steps b) and d) are carried out in the same fractionation unit and where all of the kerosene resulting from the fractionation is treated in the hydroisomerization step e). Any purging is done by withdrawing part of the residual heavy fraction.

[0162] [FIG. 1 bis] Figure 1 bis illustrates embodiment No. 1 bis in which steps b) and d) are carried out in the same fractionation unit and where all of the kerosene resulting from the fractionation is treated in the hydroisomerization step e). Any purging is done by withdrawing part of the heavy fraction resulting from said fractionation unit.

[0163] [FIG. 2] Figure 2 illustrates embodiment No. 2 in which steps b) and d) are carried out in the same fractionation unit and where only the so-called heavy kerosene fraction resulting from the fractionation is treated in the hydroisomerization step e), the so-called light kerosene fraction also resulting from said fractionation is directly mixed with the kerosene resulting from the fractionation step h) without going through step e).

[0164] [FIG. 3] Figure 3 illustrates embodiment No. 3 in which fractionation steps b) and d) are separate and where all of the kerosene from step d) is treated in hydroisomerization step e).

[0165] [FIG. 4] Figure 4 illustrates embodiment No. 4 in which fractionation steps b) and d) are separate and where only the so-called heavy kerosene fraction from step d) is treated in hydroisomerization step e), the so-called light kerosene fraction also from step d) is directly mixed with the kerosene from fractionation step h) without going through step e). [FIG. 5] Figure 5 illustrates an embodiment according to the prior art illustrated in Examples 5 and 6 in which there is no fractionation after hydrotreatment step a) and where there is only one reaction step c) in which the hydrocracking of the entire effluent from hydrotreatment step a) is carried out. In this case, the fractionation step h) is unique and allows on the one hand to separate the kerosene cut from the other cuts and on the other hand to separate the residual heavy cut which is then recycled in step c).

[0166] [FIG. 6] Figure 6 illustrates an embodiment according to the prior art illustrated in Example 8 in which after the fractionation step b) following the hydrotreatment step a), the heavy fraction is treated in a hydrocracking step c) and where the kerosene fraction from step b) is treated in a hydroisomerization step e). In this case, the fractionation step h) treats a mixture of the effluents from steps c) and e) and makes it possible, on the one hand, to separate the kerosene cut from the other cuts and, on the other hand, to separate the residual heavy cut which is then recycled in step c). Consequently, the effluent from step c) is never treated in the hydroisomerization step e).

[0167] In Figure 1, the synthesis gas 1, composed mainly of carbon monoxide and hydrogen, is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of the unit (A'), the effluent is divided into two streams: the hydrocarbon fraction 2 which includes the cold condensate and the heavy fraction corresponding to the waxes, and the non-cold condensable gas fraction 3 which can be treated directly or after possible reprocessing in the Fisher-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A) whose hydrocarbon effluent 4 is fractionated in the fractionation unit (B / D) into a gas cut 5, a naphtha cut 6, a kerosene fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which an effluent 80 is produced, the latter being able to be fractionated in a unit (G) to eliminate the purge 14, the remainder of the stream 9, purge deducted, is then injected jointly with the cut 4 resulting from hydrotreatment in the fractionation unit (B / D), the hydrocarbon streams 4 and 9 forming the stream 49. The kerosene hydrocarbon cut 7 produced in the fractionation unit (B / D) is injected into a hydroisomerization unit (E) from which the hydrocarbon effluent 10 is then separated in a fractionation unit (H). At the end of the fractionation in the unit (H), three streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12 and finally a last gas stream 13.

[0168] In Figure 1 bis, the synthesis gas 1, composed mainly of carbon monoxide and hydrogen, is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of the unit (A'), the effluent is divided into two streams: the hydrocarbon fraction 2, which includes the cold condensate and the heavy fraction corresponding to the waxes, and the non-cold condensable gas fraction 3, which can be treated directly or after possible reprocessing in the Fisher-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A), the hydrocarbon effluent 4 of which is fractionated in the fractionation unit (B / D) into a gas cut 5, a naphtha cut 6, a kerosene fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which a partially converted effluent 9 is produced, the latter then being injected jointly with the cut 4 resulting from hydrotreatment in the fractionation stage (B), the hydrocarbon streams 4 and 9 forming the stream 49. The kerosene hydrocarbon cut 7 produced by the fractionation unit (B / D) is injected into a hydroisomerization unit (E) from which the hydrocarbon effluent 10 is then separated in a fractionation unit (H). At the end of the fractionation in the unit (H), three streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12 and finally a last gas stream 13. In this embodiment, a simple withdrawal of a part of the stream 8 is possible before injection into the hydrocracking unit (C), which makes it possible to carry out a purge of a stream 14.

[0169] In Figure 2, the synthesis gas 1 is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of the unit (A'), the effluent is divided into two streams: the hydrocarbon fraction 2 which includes the cold condensate and the heavy fraction corresponding to the waxes, and the non-cold condensable gas fraction 3 which can be treated directly or after possible reprocessing in the Fischer-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A) whose hydrocarbon effluent 4 is fractionated in the fractionation unit (B / D) into a gas cut 5, a naphtha cut 6, a light kerosene fraction 71, a heavy kerosene fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which an effluent 80 is produced, the latter being able to be fractionated in a fractionation unit (G) to eliminate the purge 14, the remainder of the flow 9, purge deducted, is then injected jointly with the cut 4 resulting from hydrotreatment in the fractionation unit (B / D), the hydrocarbon flows 4 and 9 forming the flow 49. The heavy kerosene fraction 7 produced in the fractionation unit (B / D) is injected into a hydroisomerization unit (E) from which the hydrocarbon effluent 10 is then separated during in a fractionation unit (H). At the end of the fractionation in the unit (H), three streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12 and finally a last gas stream 13. The kerosene stream 11 is finally mixed during in a mixing unit (H') with the light kerosene stream 71, leading to the production of the kerosene stream 15.

[0170] In Figure 3, synthesis gas 1 is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of unit (A'), the effluent is divided into two streams: hydrocarbon fraction 2 which includes the cold condensate and the heavy fraction corresponding to the waxes, and gaseous fraction 3 which is not cold condensable and which can be treated directly or after possible reprocessing in the Fisher-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A) whose hydrocarbon effluent 4 is fractionated in the fractionation unit (B) into a gas cut 5, a naphtha cut 6, a kerosene fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which an effluent 80 is produced, the latter being able to be fractionated in a fractionation unit (G) to eliminate the purge 14, the remainder of the flow 81, purge deducted, is then injected into a fractionation unit (D) to fractionate it into a gas flow 15, a naphtha flow 16 and a kerosene flow 17. The remainder of the flow 9 corresponding to the residual heavy fraction, purge deducted, is then injected jointly with the cut 8 from the fractionation unit (B) into the hydrocracking unit (C), the hydrocarbon flows 8 and 9 forming the flow 89. The kerosene hydrocarbon cut 7 produced by the fractionation unit (B) is mixed with the cut 17 from the fractionation unit (D) to form the flow 710, the latter then being injected into a hydroisomerization unit (E) from which the hydrocarbon effluent 10 is then separated in a fractionation unit (H).At the end of the fractionation in unit (H), three streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12 and finally a last gas stream 13.

[0171] In Figure 4, synthesis gas 1 is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of the unit (A'), the effluent is divided into two streams: hydrocarbon fraction 2 which includes the cold condensate and the heavy fraction corresponding to the waxes, and the non-cold condensable gas fraction 3 which can be treated directly or after possible reprocessing in the Fischer-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A) whose hydrocarbon effluent 4 is fractionated in a fractionation unit (B) into a gas cut 5, a naphtha cut 6, a kerosene fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which an effluent 80 is produced, the latter being able to be fractionated in a fractionation unit (G) to eliminate the purge 14, the remainder of the flow 81, purge deducted, is then injected into the fractionation unit (D) to fractionate it into a gas flow 50 and a naphtha flow 60 and at least one heavy kerosene flow 71 and a light kerosene flow 72. The remainder of the flow 9 corresponding to the residual heavy fraction purge deducted is then injected jointly with the heavy fraction 8 from the fractionation unit (B) into the hydrocracking unit (C), the hydrocarbon flows 8 and 9 forming the flow 89. The kerosene hydrocarbon cuts 7 and 71 produced respectively by the fractionation units (B) and (D) form the flow 710 which is then injected in a hydroisomerization unit (E) from which the hydrocarbon effluent 10 is then separated in a fractionation unit (H).At the end of the fractionation in the unit (H), three streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12 and finally a last gas stream 13. The kerosene stream 11 is finally mixed in a mixing unit (H') with the light kerosene stream 72, leading to the production of the kerosene stream 15.

[0172] In Figure 5, synthesis gas 1 is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of unit (A'), the effluent is divided into two streams: hydrocarbon fraction 2 which includes the cold condensate and the heavy fraction corresponding to the waxes, and gaseous fraction 3 which is not cold condensable and which can be treated directly or after possible reprocessing in the Fischer-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A) whose hydrocarbon effluent 4 is then injected into a hydrocracking unit (C) whose hydrocarbon effluent 6 is then separated in a fractionation unit (H). At the end of the fractionation in unit (H), three streams are produced: a stream corresponding to the kerosene cut 9, a naphtha stream 8 and finally a last gas stream 7.The residual heavy fraction 10 is injected jointly with the hydrocarbon effluent 4 from the hydrotreatment unit (A) into the hydrocracking unit (C), the hydrocarbon streams 10 and 4 forming the stream 104.

[0173] In Figure 6, synthesis gas 1 is directed to a Fischer-Tropsch synthesis unit (A'). At the outlet of unit (A'), the effluent is divided into two streams: hydrocarbon fraction 2 which includes the cold condensate and the heavy fraction corresponding to the waxes, and non-cold condensable gas fraction 3 which can be treated directly or after possible reprocessing in the Fischer-Tropsch synthesis unit (recycle). Fraction 2 is sent to the hydrotreatment unit (A) whose hydrocarbon effluent 4 is fractionated in a fractionation unit (B) into a gas cut 5, a naphtha cut 6, a kerosene fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which an effluent 10 is produced. The kerosene fraction 7 produced by the fractionation unit (B) is injected into a hydroisomerization unit (E) whose hydrocarbon effluent 9 is then mixed with the effluent 10 to form the stream 90, the latter being separated in a fractionation unit (H). Following the fractionation in the unit (H), three streams are produced: a stream corresponding to the kerosene cut 13, a naphtha stream 12 and finally a last gas stream 11. The residual heavy fraction 91 is recycled into the hydrocracking unit (C) mixed with the heavy fraction (8) to form the stream (918).

[0174] The examples illustrate the invention without limiting its scope.

[0175] Example 1: Preparation of hydrotreatment catalyst C1 (compliant).

[0176] The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with molybdenum oxide MoOs 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.

[0177] Example 2: Preparation of hydrocracking catalyst C2 (compliant).

[0178] 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 SiO2 in the final solid. This mixture is quickly homogenized in a commercial colloidal mill in the presence of nitric acid so that the nitric acid content by weight 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 provided 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.

[0179] The characteristics of the support thus prepared are as follows:

[0180] - an average mesopore diameter measured by mercury porosimetry of 8.0 nm,

[0181] - a total pore volume measured by mercury porosimetry of 0.49 ml / g,

[0182] - a mesoporous volume measured by mercury porosimetry of 0.48 ml / g,

[0183] - a volume of macropores, the diameter of which is greater than 50 nm less than 0.02 ml / g,

[0184] - a BET surface area of ​​235 m 2 / g. 50 grams of silica-alumina extrudates are then subjected to a dry impregnation step with an aqueous solution of platinum tetramine nitrate, left to mature in a water 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 platinum content by weight of the finished catalyst after calcination is 0.16%, its dispersion measured by H2 / O2 titration is 33%, its distribution coefficient measured by Castaing microprobe is 0.96.

[0185] Example 3: Preparation of a C3 hydroisomerization catalyst (compliant).

[0186] Synthesis of IZM-2 zeolite.

[0187] 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), 1,6bis(methylpiperidinium)hexane dibromide structuring agent, aluminum hydroxide (Aldrich) and deionized water. The molar composition of the mixture is as follows: 1 SiO2; 0.0060 AI2O3; 0.1666 Na2 <D; 0,1666 1 ,6bis(méthylpiperidinium)hexane; 33,3333 H2O. Le mélange est agité vigoureusement pendant une demi-heure. Le mélange est ensuite transféré, après homogénéisation, dans un autoclave de type PARR. L’autoclave est chauffé pendant 5 jours à 170°C sous agitation en tourne broche (30 tours / min). Le produit obtenu est filtré, lavé à l’eau déionisée pour atteindre un pH neutre puis séché une nuit à 100°C en étuve. Le solide est ensuite introduit dans un four à moufle pour y être calciné afin d’éliminer le structurant.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 traversed 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:

[0188] - weight percentage of hexacoordinated aluminum atoms AI VI : 5%,

[0189] - ratio of the number of moles of silicon divided by the number of moles of network aluminum, in mole / mole, Si / Al: 72, - ratio of the number of moles of sodium divided by the number of moles of network aluminum, in mole / mole, Na / Al: 0.03.

[0190] Preparation of IZM-2 / alumina support.

[0191] 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.

[0192] Impregnation of platinum on IZM-2 / alumina support.

[0193] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)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:

[0194] - increase in temperature from ambient to 450°C at 5°C / min;

[0195] - two-hour stage at 450°C;

[0196] - descent to ambient.

[0197] 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.

[0198] Example 4: Preparation of a C4 hydroisomerization catalyst (compliant).

[0199] Zeolite ZSM-12.

[0200] 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 access to the following results for ZSM-12:

[0201] - 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 the ZSM-12 / alumina support. 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 quadrilobe 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.

[0205] Impregnation of platinum on ZSM-12 / alumina support.

[0206] Platinum impregnation is carried out by dry impregnation of the support with an aqueous solution containing platinum tetramine nitrate Pt(NH3)4(NO3)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:

[0207] - increase in temperature from ambient to 450°C at 5°C / min;

[0208] - two-hour stage at 450°C;

[0209] - descent to ambient.

[0210] 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.

[0211] Example 5: Process for producing kerosene not in accordance with the invention.

[0212] In this example illustrated in Figure 5, there is no step of fractionation of the feed into at least one gaseous fraction, one naphtha fraction, one kerosene fraction and one heavy fraction after hydrotreatment step nor isomerization step of the kerosene fraction.

[0213] A paraffinic feedstock from Fischer-Tropsch synthesis has the characteristics given in Table 1. This feedstock includes cold condensates and waxes.

[0214] [Table 1]

[0215] Table 1: Characteristics of the charge resulting from Fischer-Tropsch synthesis

[0216] The total atomic oxygen content in said fraction is measured by the infrared absorption technique described in patent application US2009 / 0018374A1.

[0217] The oxygen content represents the content of molecules comprising at least one oxygen atom present in the heavy fraction and is expressed as a percentage by weight relative to the total pass of said fraction. The oxygen content is measured by gas chromatography.

[0218] Step a) hydrotreatment of the paraffinic feedstock.

[0219] The paraffinic feedstock is treated in a lost hydrogen flow-through bed, i.e. without hydrogen recycling, on the C1 hydrotreatment catalyst under operating conditions which allow the elimination of olefinic and oxygenated compounds as well as traces of nitrogen.

[0220] The selected operating conditions are as follows: hourly volumetric flow rate WH (charge volume / catalyst volume / hour) = 2 h 1 , total working pressure: 7 MPa, hydrogen / charge ratio: 700 normal liters / liter, temperature: 330°C.

[0221] Before testing, the C1 catalyst undergoes a reduction step under the following operating conditions: pure hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst, rise from ambient temperature to 120°C: 10°C / min, one-hour hold at 120°C, rise from 120°C to 450°C at 5°C / min, two-hour hold at 450°C, total pressure: 7 MPa The hydrotreated effluent then undergoes a flash separation and decantation step. At the end of said separation, carbon monoxide and / or carbon dioxide and / or water and / or ammonia and / or hydrocarbons with less than five atoms possibly formed during the hydrotreatment are eliminated. The olefin and nitrogen compound contents of the hydrotreated effluent having undergone said separation step fall below the detection thresholds and the oxygen content measured by IR is less than 500 ppm by weight, while the conversion of the 300°C fraction + in fraction 300 00 is negligible (less than 5% by weight). The characteristics of the heavy fraction after hydrotreatment and separation are shown in Table 2.

[0222] [Table 2]

[0223] Table 2: Characteristics of the effluent after hydrotreatment and separation.

[0224] Step c) hydrocracking of the hydrotreated and separated paraffinic feedstock.

[0225] The hydrotreated paraffinic feedstock having undergone the flash and decantation step mixed with the residual heavy fraction from step c) below constitutes the hydrocracking feedstock sent to the hydrocracking catalyst C2 in accordance with the invention.

[0226] Before testing, the C2 catalyst undergoes a reduction step under the following operating conditions: pure hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst, rise from ambient temperature to 120°C: 10°C / min, one-hour hold at 120°C, rise from 120°C to 450°C at 5°C / min, two-hour hold at 450°C, pressure: 6.5 MPa

[0227] After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to charge ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 1 h -1 The hydrocracking feedstock is treated in a lost hydrogen through-bed, i.e. without hydrogen recycling. The conversion of the 300°C fraction + is taken equal to:

[0228] C(300°C + ) = [ (% of 300°C- effluents ) - (% of 300°C- load) ] / [ 100 - (% of 300°C- load)] with

[0229] % of 300°C effluents: mass percentage of compounds with boiling points below 300°C in hydrocracking effluents, and

[0230] % of 300°C _ charge: mass percentage of compounds having boiling points below 300°C in the hydrocracking charge.

[0231] The reaction temperature is adjusted to 325°C so as to obtain a conversion level of the 300°C fraction + equal to 0.35. This corresponds to a recycle rate equal to 2.85. The conversion level of the 300°C fraction + was adjusted to obtain a value of -40°C for the crystal disappearance temperature for the kerosene obtained after step h) of fractionation of the hydrocracking effluent.

[0232] Step h) fractionation of the hydrocracking effluent.

[0233] The hydrocracking effluent obtained from step b) is fractionated by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C, and a residual heavy fraction with an initial boiling point above 300°C. Said residual heavy fraction is then recycled to the hydrocracking step b). The disappearance point of the crystals of the kerosene cut is then measured according to standard ASTM D5972. Material balances carried out over 24 hours allow the yields to be calculated in different cuts: yield in light gases: mass of gas Ci to C4 / mass of paraffinic charge * 100, yield in naptha cut: mass of liquid effluent with initial boiling point of 35°C and final boiling point of 120°C / mass of paraffinic charge * 100,kerosene cut yield: mass of liquid effluent with initial boiling point of 120°C and final boiling point of 300°C / mass of paraffinic feed * 100, or the mass of gas Ci to C4 corresponds to the quantity of gas Ci to C4 generated during 24 hours at the end of step a) and step c); the mass of paraffinic feed corresponds to the quantity of paraffinic feed from Fischer-Tropsch synthesis consumed during 24 hours,the mass of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C corresponds to the quantity of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C generated during 24 hours at the end of step c) and the mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C corresponds to the quantity of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C generated during 24 hours at the end of step c). Table 3 reports the yields in different cuts as well as the properties of the kerosene cut obtained.,

[0234] Example 6: Process for producing kerosene not in accordance with the invention illustrated in Figure 5.

[0235] The same paraffinic feedstock from Fischer-Tropsch synthesis is used as in Example 5. Example 6 is illustrated in Figure 5.

[0236] Step a) hydrotreatment of the paraffinic feedstock.

[0237] Same as example 5.

[0238] Step c) hydrocracking of the hydrotreated and separated paraffinic feedstock.

[0239] The hydrotreated paraffinic feedstock having undergone the flash and decantation step mixed with the residual heavy fraction from step c) below constitutes the hydrocracking feedstock sent to the C4 hydroisomerization catalyst in accordance with the invention.

[0240] Before testing, the C4 catalyst undergoes a reduction step under the following operating conditions: pure hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst, rise from ambient temperature to 120°C: 10°C / min, one-hour hold at 120°C, rise from 120°C to 450°C at 5°C / min, two-hour hold at 450°C, pressure: 6.5 MPa.

[0241] After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to charge ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 2 h-1 .

[0242] The hydrocracking feedstock is treated in a lost hydrogen through-bed, i.e. without hydrogen recycling. The conversion of the 300°C+ fraction is taken as:

[0243] C(300°C + ) = [ (% of 300°C- effluents ) - (% of 300°C- load) ] / [ 100 - (% of 300°C- load)] with

[0244] % of 300°C _ effluents: mass percentage of compounds with boiling points below 300°C in the effluents, and

[0245] % of 300°C' feed: mass percentage of compounds with boiling points below 300°C in the hydrocracking feed. The reaction temperature is adjusted to 320°C so as to obtain a conversion level of the 300°C fraction + equal to 0.35. This corresponds to a recycle rate equal to 2.85. The conversion level of the 300°C fraction +was adjusted to obtain a value of -40°C for the crystal disappearance temperature for the kerosene obtained after step h) of distillation of the hydrocracking effluent.

[0246] Step h) fractionation of the hydrocracking effluent.

[0247] Same as example 5.

[0248] Table 3 shows the yields in different cuts as well as the properties of the kerosene cut obtained.

[0249] Example 7: Process for producing kerosene not in accordance with the invention.

[0250] The same paraffinic feedstock from Fischer-Tropsch synthesis is used as in Example 5. The example is illustrated in Figure 6.

[0251] Step a) hydrotreatment of the paraffinic feedstock.

[0252] Same as example 5.

[0253] Step b) fractionation of the hydrotreated and separated paraffinic feedstock.

[0254] The hydrotreated paraffinic feedstock having undergone the flash and decantation step is fractionated by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C, and a heavy fraction with an initial boiling point above 300°C.

[0255] Step e) hydroisomerization of the kerosene fraction from step b).

[0256] The kerosene fraction from step b) constitutes the hydroisomerization feedstock sent to the C4 hydroisomerization catalyst according to the invention. The reduction conditions of the C4 catalyst are identical to those of Example 6. The hydroisomerization feedstock is treated in a lost hydrogen through-bed, i.e. without hydrogen recycling.

[0257] After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to charge ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 2 h -1 .

[0258] The reactor temperature is adjusted to achieve a kerosene yield of 95%. The kerosene yield is calculated as (% of 120-300°C effluents) / (% of 120-300°C feed) * 100 with: % of 120-300°C effluents: mass percentage of compounds with boiling points between 120 and 300°C in the effluents leaving the hydroisomerization unit;

[0259] % of 120-300°C feed: mass percentage of compounds with boiling points between 120 and 300°C in the feed entering the hydroisomerization unit;

[0260] Step c) hydrocracking of the heavy fraction from step b).

[0261] The heavy fraction from step b) mixed with the residual heavy fraction from step h) below constitutes the hydrocracking feed sent to the hydrocracking catalyst C2 in accordance with the invention. The reduction conditions of catalyst C2 are identical to those of example 5.

[0262] After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to charge ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 1 h 1 .

[0263] The hydrocracking feedstock is treated in a lost hydrogen through-bed, i.e. without hydrogen recycling. The conversion of the 300°C+ fraction is taken as:

[0264] C(300°C + ) = [ (% of 300°C- effluents ) - (% of 300°C- load) ] / [ 100 - (% of 300°C- load)] with

[0265] % of 300°C effluents: mass percentage of compounds with boiling points below 300°C in the effluents, and

[0266] % of 300°C- feed: mass percentage of compounds having boiling points below 300°C in the hydrocracking feed.

[0267] The reaction temperature is adjusted to 340°C so as to obtain a conversion level of the 300°C fraction + equal to 0.45. This corresponds to a recycle rate equal to 2.22. The conversion level of the 300°C fraction + was adjusted to obtain a crystal disappearance temperature value for the kerosene obtained after fractionation step h) equal to -40°C measured according to ASTM D5972.

[0268] Step h) fractionation of the hydrocracking effluent and the hydroisomerization effluent.

[0269] The hydrocracking effluent obtained at the end of step c) is fractionated at the same time as the effluent from the hydroisomerization step e) by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C, and a residual heavy fraction with an initial boiling point above 300°C. Said residual heavy fraction is then recycled to the hydrocracking step c). The fractionation of the hydrocracking effluent is carried out in a distillation column separate from that used in step b). There is therefore no joint fractionation of the hydrocracking effluent and the hydrotreated and separated paraffinic feedstock.

[0270] Material balances carried out over 24 hours make it possible to calculate the yields in different cuts: light gas yield: mass of Ci to C4 gas / mass of paraffinic feedstock * 100, naptha cut yield: mass of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C / mass of paraffinic feedstock * 100, kerosene cut yield: mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C / mass of paraffinic feedstock * 100, or the mass of Ci to C4 gas corresponds to the quantity of Ci to C4 gas generated during 24 hours at the end of step a) of hydrotreatment of the paraffin feedstock from the Fischer-Tropsch synthesis and step h) of joint fractionation of the hydrocracking and hydroisomerization effluents;the mass of paraffinic feed corresponds to the quantity of paraffinic feed from the Fischer-Tropsch synthesis consumed during 24 hours, the mass of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C corresponds to the quantity of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C generated during 24 hours at the end of step b) of fractionation after hydrotreatment of the paraffin feed from the Fischer-Tropsch synthesis and step h) of joint fractionation of the hydrocracking and hydroisomerization effluents and the mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C corresponds to the quantity of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C generated during 24 hours at the end of step h) of joint fractionation of the hydrocracking effluents and hydroisomerization. The disappearance point of the crystals of said kerosene cut is then measured according to the ASTM D5972 standard.;

[0271] Table 3 shows the yields in different cuts as well as the properties of the kerosene cut obtained.

[0272] Example 8: Process for producing kerosene not in accordance with the invention.

[0273] The same paraffinic feedstock from Fischer-Tropsch synthesis is used as in Example 5.

[0274] Example 8 is illustrated in Figure 1 in the case where the purge is zero.

[0275] Step a) hydrotreatment of the paraffinic feedstock.

[0276] Same as example 5.

[0277] Step b) fractionation of the hydrotreated and separated paraffinic feedstock.

[0278] Same as example 7.

[0279] Step c) hydrocracking of the heavy fraction from step b). The heavy fraction from step b) mixed with the residual heavy fraction from step d) below constitutes the hydrocracking feed sent to the hydrocracking catalyst C2 according to the invention. The reduction conditions of catalyst C2 are identical to those of Example 5. After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to feed ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 1 h -1 .

[0280] The hydrocracking feedstock is treated in a lost hydrogen through-bed, i.e. without hydrogen recycling. Conversion of the 300°C fraction + is taken equal to:

[0281] C(300°C + ) = [ (% of 300°C- effluents ) - (% of 300°C- load) ] / [ 100 - (% of 300°C- load)] with

[0282] % of 300°C' effluents = mass percentage of compounds having boiling points lower than

[0283] 300°C in the effluents, and

[0284] % of 300°C _ charge = mass percentage of compounds having boiling points below 300°C in the hydrocracking charge.

[0285] The reaction temperature is adjusted to 350°C so as to obtain a conversion level of the 300°C fraction + equal to 0.65 not in accordance with the invention. This corresponds to a recycling rate equal to 1.5.

[0286] Step d) fractionation of the hydrocracking effluent.

[0287] The hydrocracking effluent obtained from step c) is fractionated by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C, and a residual heavy fraction with an initial boiling point above 300°C. Said residual heavy fraction is then recycled to hydrocracking step c). The fractionation of the hydrocracking effluent is carried out in the same distillation column as that used in step b). There is therefore a joint fractionation of the hydrocracking effluent and the hydrotreated and separated paraffinic feedstock. of the kerosene fraction from the

[0288] The kerosene fraction from step b) constitutes the hydroisomerization feedstock sent to the C4 hydroisomerization catalyst according to the invention. The reduction conditions of the C4 catalyst are identical to those of Example 7. The hydroisomerization feedstock is treated in a lost hydrogen cross-bed, i.e. without recycling hydrogen.

[0289] After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to charge ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 2 h -1 .

[0290] The temperature is adjusted so as to obtain a value of -40°C for the crystal disappearance temperature for the kerosene obtained after the fractionation step h).

[0291] Step h) fractionation of the hydroisomerization effluent.

[0292] The hydroisomerization effluent obtained at the end of step e) is fractionated by atmospheric distillation into different hydrocarbon cuts, namely light gases from C1 to C4, a naphtha cut with an initial boiling point of 35°C and a final boiling point of 120°C, a kerosene cut with an initial boiling point of 120°C and a final boiling point of 300°C. This fractionation is carried out in an atmospheric distillation unit different from step d).

[0293] Material balances carried out over 24 hours make it possible to calculate the yields in different cuts: yield in light gases: mass of gas Ci to C4 / mass of paraffinic feedstock * 100, yield in naphtha cut: mass of liquid effluent with initial boiling point of 35°C and final boiling point of 120°C / mass of paraffinic feedstock * 100, yield in kerosene cut: mass of liquid effluent with initial boiling point of 120°C and final boiling point of 300°C / mass of paraffinic feedstock * 100, or the mass of gas Ci to C4 corresponds to the quantity of gas Ci to C4 generated during 24 hours at the end of step a) of hydrotreatment of the paraffin feedstock from the Fischer-Tropsch synthesis,from step b) and d) of joint fractionation of the hydrotreatment effluent and separation of the paraffin feedstock from the Fischer-Tropsch synthesis and the hydroconversion effluent and from step h) of fractionation of the hydroisomerization effluent; the mass of paraffinic feedstock corresponds to the quantity of paraffinic feedstock from the Fischer-Tropsch synthesis consumed over 24 hours,the mass of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C corresponds to the quantity of liquid effluent with an initial boiling point of 35°C and a final boiling point of 120°C generated over 24 hours at the end of steps b) and d) of joint fractionation of the hydrotreatment effluent and separation of the paraffinic feedstock from the Fischer-Tropsch synthesis and the hydrocracking effluent and step h) of fractionation of the hydroisomerization effluent and the mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C corresponds to the quantity of liquid effluent with an initial boiling point of 120°C and a final boiling point of 300°C generated over 24 hours at the end of step h) of fractionation of the hydroisomerization effluent. The disappearance point of the crystals of said kerosene cut is then measured according to the ASTM D5972 standard.

[0294] Table 3 shows the yields in different cuts as well as the properties of the kerosene cut obtained. Example 9: Process for producing kerosene in accordance with the invention.

[0295] Example 9 is illustrated in Figure 1 in the case where the purge is zero. of the paraffinic charge.

[0296] Same as example 8. b) load splitting bedded and seated

[0297] Same as example 8. of the heavy fraction from the

[0298] Same as Example 8, except that the reaction temperature is adjusted to 330°C so as to obtain a conversion level of the 300°C fraction + equal to 0.2. This corresponds to a recycle rate equal to 5. fractionation of the effluent of

[0299] Same as example 8. of the kerosene fraction from the

[0300] Same as example 8. fractionation of the effluent of

[0301] Same as example 8.

[0302] The yields in different cuts are calculated identically to example 8, the disappearance point of the crystals of the kerosene cut is also measured according to the ASTM D5972 standard.

[0303] Table 3 shows the yields in different cuts as well as the properties of the kerosene cut obtained.

[0304] Example 10: Process for producing kerosene in accordance with the invention.

[0305] Example 10 is illustrated in Figure 1 in the case where the purge is zero. EU.

[0306] Same as example 9. b) load splitting bedded and seated

[0307] Same as example 9. c) hydrocracking of the heavy fraction from step b).

[0308] Same as example 9. fractionation of the effluent of

[0309] Same as example 9.

[0310] Step e) hydroisomerization of the kerosene fraction from step b).

[0311] The kerosene fraction from step b) constitutes the hydroisomerization feedstock sent to the hydroisomerization catalyst C3 according to the invention. The reduction conditions of catalyst C3 are identical to those of catalyst C4 of example 6. The hydroisomerization feedstock is treated in a lost hydrogen cross-bed, i.e. without recycling hydrogen.

[0312] After reduction, the catalytic test is carried out under the following conditions: total pressure of 6.5 MPa, hydrogen to charge ratio of 800 normal liters / liter, hourly volumetric flow rate (WH) equal to 2 h -1 .

[0313] The temperature is adjusted so as to obtain a value of -40°C for the crystal disappearance temperature for the kerosene obtained after the fractionation step h) fractionation of the effluent from

[0314] Same as example 9.

[0315] The yields in different cuts are calculated identically to example 9, the disappearance point of the kerosene cut crystals is also measured according to ASTM D2386.

[0316] Table 3 shows the yields in different cuts as well as the properties of the kerosene cut obtained.

[0317] [Table 3]

[0318] Table 3: Yields in different cuts and properties of the kerosene cuts obtained for the different examples of the invention. The examples provided highlight the importance of carrying out the process according to the invention. In particular, the comparison of Example 9 according to the invention with Examples 5 and 6 demonstrates that to obtain a kerosene whose crystal disappearance point is less than or equal to the standard, it is necessary to operate with two types of catalysts located in two different reactors. Furthermore, the comparison of Example 9 according to the invention with Example 7 demonstrates that it is necessary to arrange the hydroisomerization step in series after the hydrocracking step. Indeed, when the two steps are operated in parallel, as is the case in Example 7, the kerosene yield is lower. Finally, the comparison of Example 9 according to the invention with the comparative Example 8 highlights the importance of placing the hydrocracking reactor in step c) at low conversion since otherwise the kerosene yield is also lower.Finally, examples 9 and 10 demonstrate that the process according to the invention using two types of hydroisomerization catalysts according to the invention makes it possible to maximize the kerosene yield.

Claims

CLAIMS 1. A process for producing kerosene from a paraffinic feedstock produced by Fischer-Tropsch synthesis comprising at least the following steps and preferably consists of: a) Hydrotreatment of at least a portion and preferably all of the paraffinic feedstock produced by Fischer-Tropsch synthesis in the presence of a hydrotreatment catalyst and operating at a temperature of between 250 and 450°C, at a pressure of between 0.5 and 15 MPa, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 3000 normal liters of hydrogen per liter of feedstock, and at an hourly volumetric flow rate of between 0.1 and 40 h 1, b) Fractionation in one or more stages of the effluent from stage a) into at least one gaseous fraction, a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 100 and 160°C, at least one kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C, and a heavy fraction having an initial boiling point above 250°C, c) Hydrocracking all of said heavy fraction having an initial boiling point above 250°C from stage b) to produce a hydrocracked effluent, in the presence of a hydrocracking catalyst and operating at a temperature of between 250 and 450°C, at a pressure of between 0.2 and 15 MPa, at a rate spatial between 0.1 and 10 h 1, and at a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed, said hydrocracking catalyst comprising at least one noble metal from group VIII of the periodic table and a support comprising at least one acidic solid, said operating conditions of said step c) being adjusted so that the conversion per pass into products having boiling points greater than or equal to a temperature Tx into products with boiling points lower than this same temperature Tx is less than 0.5, preferably less than 0.4 and preferably less than 0.3 and more preferably less than 0.2, Tx being between 250 and 320°C, preferably between 270 and 310°C and very preferably between 280 and 300°C, d) Fractionation in one or more stages of at least all or part of the effluent from step c), into at least one fraction gaseous,a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 100 and 160°C, at least one kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C, and a residual heavy fraction having an initial boiling point above 250°C, said fractionation steps b) and d) being advantageously able to be carried out in the same units or in separate fractionation units, e) The hydroisomerization of at least a portion of the kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C from fractionation step b) and of all or part of the kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of between 250 and 320°C from fractionation step d) in the case where said fractionation steps b) and d) are distinct, in the presence of a hydroisomerization catalyst operating at a temperature of between 200 and 450°C, a pressure of between 1 and 15 MPa, a space velocity of between 0.1 and 10 h 1, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of feed, to produce a hydroisomerized effluent, said hydroisomerization catalyst comprising at least one noble metal from group VIII of the periodic table and a support comprising at least one binder and at least one zeolite chosen from zeolites of structural type TON, FER, EUO, AEL, *MRE, MTW, MOR, *BEA and zeolite IZM-2, taken alone or as a mixture, f) Recycling all or part of the residual heavy fraction having an initial boiling point above 250°C in hydrocracking step c), g) optionally purging part of the hydrocracked effluent from hydrocracking step c), h) Fractionating all or part of the hydroisomerized effluent from step e) hydroisomerization, into at least one gaseous fraction,a naphtha fraction having an initial boiling point of between 30 and 60°C and a final boiling point of between 120 and 160°C, and at least one kerosene fraction having an initial boiling point of between 120 and 160°C., 2. Process according to claim 1, in which two kerosene fractions are separated during fractionation step b), a light kerosene fraction whose initial boiling point is between 100 and 160°C and the final boiling point is between 160 and 220°C, and a heavy kerosene fraction whose initial boiling point is between 160 and 220°C and the final boiling point is between 250 and 320°C, and only the heavy kerosene fraction is treated in said hydroisomerization step e) and the light kerosene fraction from fractionation step b) can then advantageously be mixed with the kerosene fraction from fractionation step h).

3. Process according to one of claims 1 or 2 in which the hydrocracking catalyst used in hydrocracking step c) comprises at least one noble metal, said noble metal being platinum and a silica-alumina as Bronsted acid solid, without any other binder.

4. Process according to one of the preceding claims, in which, in the case where fractionation steps b) and d) are carried out in separate fractionation units, two kerosene fractions are separated during fractionation step d), a light kerosene fraction having an initial boiling point of between 100 and 160°C and a final boiling point of less than or equal to 250°C, preferably less than 230°C and very preferably less than 200°C, and a heavy kerosene fraction having an initial boiling point of more than 200°C and a final boiling point of between 250 and 320°C, and only the heavy kerosene fraction from fractionation step d) is sent to said hydroisomerization step e), mixed with the kerosene fraction from fractionation step b).

5. Process according to one of claims 1 to 4 in which the catalyst of step e) comprises and is preferably constituted by platinum, and a support comprising and preferably constituted by a ZSM-12 zeolite and an alumina binder.

6. Process according to one of claims 1 to 4 in which the catalyst of step e) comprises and is preferably constituted by platinum, and a support comprising and preferably constituted by an IZM-2 zeolite and an alumina binder.

7. Process according to one of claims 1 to 6 in which in the variant or steps b) and d) are carried out in a single fractionation unit, the hydrocracked effluent from step c), optionally purged, is advantageously recycled in said common fractionation step in a mixture with all or part of the effluent from hydrotreatment step a).

8. Method according to one of claims 1 to 7 in which in the variant or fractionation steps b) and d) are carried out in the same fractionation unit, and in the case where a light kerosene fraction comes from said fractionation step, said light kerosene fraction is mixed with the kerosene fraction coming from the final fractionation step h).

9. Method according to one of claims 1 to 7 in which in the variant or fractionation steps b) and d) are carried out in separate fractionation units, and in the case where one or more light kerosene fraction(s) is (are) derived from said fractionation steps b) and / or d), said light kerosene fraction(s) is (are) mixed with the kerosene fraction resulting from the final fractionation step h).

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