Improved process for producing middle distillates from serially synthesized hydrocarbons with low conversion per pass

Abstract

An objective of the present invention is therefore to provide a process for producing middle distillates and in particular diesel 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 distillate cut, preferably diesel, 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. Figure 1 to be published

Description

Title of the invention: Improved process for producing middle distillates from low conversion serially synthesized hydrocarbons per pass Field of invention

[0001] The Fischer-Tropsch process makes it possible to produce synthetic hydrocarbons from a gaseous feedstock composed mainly of hydrogen and carbon monoxide, also called synthesis gas (CO+H2).

[0002] In particular, in the low-temperature Fischer-Tropsch process, the synthesis gas is catalytically transformed into water, oxygenated products, olefins and essentially linear paraffins also called n-paraffins, in gaseous, liquid or solid form under standard conditions. After separation of the water, the synthetic hydrocarbons obtained are mainly composed of predominantly linear paraffins but also contain olefins and oxygenated products.

[0003] 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.

[0004] 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 18 carbon atoms and a boiling point of 316°C is approximately 26°C, which makes its direct incorporation into the diesel cut impossible given the required specification of -5°C or +5°C depending on the season for the cloud point for diesel prepared from paraffins produced by Fischer-Tropsch synthesis according to standard EN15940. A hydroisomerization step is necessary to lower the cloud points of the various hydrocarbon compounds. Furthermore, treatment steps are necessary to remove olefinic and oxygenated compounds, increase the yield of middle distillates and improve the properties of the cuts produced 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 beyond 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 enter the interval. distillation desired. For the light fraction, i.e. 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, one can then choose to play on the formulation of the 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 the 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 according to the desired objective, i.e. kerosene.

[0005] In order to illustrate the problem, we will cite below some results from the prior art allowing it to be evaluated.

[0006] 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 cuts, the lighter of the two then being mixed with the cold condensate cut 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 jointly with the effluent from 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, which does not allow the problem of maximizing the diesel yield to be addressed at the desired specifications.

[0007] 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.

[0008] 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 type platinum / silica-alumina for the first and a molecular sieve and preferably of the platinum / zeolite ZSM-12 type with a silicic binder for the second.

[0009] An optional vacuum distillation step of the residual fraction may 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 may be recycled to the hydrocracking and hydroisomerization reaction step. The base oil distillate fraction may 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 silicic 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 - (approximately 20% by weight of base oil is not recycled), 10% by weight of gas is formed and only 67 to 71% by weight 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 respond to the problem posed of maximizing the diesel yield to the desired specifications.

[0010] Patent application WO05001006 also describes a method for treating a hydrocarbon wax, which may originate from 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 diesel yield to the desired specifications.

[0011] Patent application WO12001546 again covers a process for producing diesel, heavy distillate and residual base oil from a feedstock derived from FT comprising a hydroconversion / hydroisomerization step for obtaining an at least partially isomerized feedstock followed by a step for 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 reveal that with such a process, the production of distillates, in this case diesel, remains low, a significant proportion of the charge not being converted into a lighter cut and being reserved for the production of base oil, almost 50% by weight.

[0012] This is also the case of patent application WO05003067 in which a process for producing base oil from a feedstock produced by Fischer-Tropsch synthesis is described which first comprises a separation of 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 on a catalyst which 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 diesel yield to the desired specifications.

[0013] Finally, patent application WO12001552 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 the heavy residual fraction also undergoes dewaxing, the latter forming a base oil, which does not make it possible to address the problem addressed by maximizing the diesel yield to the desired specifications.

[0014] Application WO09041478 discloses a process for producing diesel fuel comprising firstly a step of fractionation of 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 by the yield since a fraction of kerosene or diesel will necessarily have to be set aside.

[0015] The research work carried out by the applicant led him to discover that in order to selectively produce a gas oil cut with the properties and specifications in force, in particular for the cutting points and the 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.

[0016] An objective of the present invention is therefore to provide a process for producing middle distillates and in particular diesel 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 middle distillate cut, preferably diesel, while minimizing the production of light cuts such as light C1-C4 gases, naphtha and the production of oil, the latter being limited to a possible purge on an effluent recycle. Subject of the invention

[0017] The present invention relates to a process for producing middle distillates and preferably diesel from a paraffinic feedstock produced by Fischer-Tropsh synthesis comprising at least the following steps and is preferably constituted by:

[0018] a) hydrotreatment of at least part 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 h1,

[0019] 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 200°C, at least one middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C, and a heavy fraction having an initial boiling point of between 350 and 390°C.

[0020] c) hydrocracking all of said heavy fraction having an initial boiling point of between 350 and 390°C from 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 h 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 350 and 390°C, preferably between 360 and 380°C,

[0021] d) Fractionation in one or more stages of at least all or part of the effluent from stage c), 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 200°C, at least one middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C, and a residual heavy fraction having an initial boiling point of between 350 and 390°C, said fractionation stages b) and d) being advantageously able to be carried out in the same fractionation unit or in separate fractionation units,

[0022] e) T hydroisomerization of at least a portion of the middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C from fractionation step b) and of all or part of the middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°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 h1, 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,

[0023] f) recycling all or part of the residual heavy fraction having an initial boiling point above 350°C in hydrocracking step c),

[0024] g) optionally purging part of the hydrocracked effluent from hydrocracking step c),

[0025] h) fractionating all or part of the hydroisomerized effluent from hydroisomerization step e) 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 200°C, and at least one middle distillate fraction having an initial boiling point of between 100 and 200°C.

[0026] 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 should first be combined before their simultaneous hydrotreatment.

[0027] Generally, said process applies to all or part of the effluent resulting from Fischer-Tropsch synthesis.

[0028] An advantage of the present invention is therefore to provide a process for producing middle distillates from a paraffinic effluent produced by Fischer-Tropsch synthesis making it possible both to maximize the production of middle distillates and to limit the production of light cracked products which cannot be incorporated into a kerosene and / or diesel pool and therefore to improve the selectivity towards the production of the latter.

[0029] Throughout the rest of the text, middle distillates are understood to mean the sum of the kerosene and diesel cuts.

[0030] 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 of between 350 and 390°C, preferably between 360 and 380°C.

[0031] 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 of between 350 and 390°C, preferably between 360 and 380°C.

[0032] 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 middle distillate fraction having an initial boiling point of between 100 and 200°C and a boiling point final boiling point between 350 and 390°C. Thus, the residual heavy fraction not recycled is reduced to the strict minimum, preferably, it is zero and, failing that, it corresponds to a purge.

[0033] In the following text, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, publisher CRC press, editor-in-chief DR Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUP AC classification, and group VIB to the metals of column 6.

[0034] 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.

[0035] In the present description, the expression "greater than..." is understood as strictly greater, and symbolized by the sign ">", and the expression "less than" as strictly less, and symbolized by the sign "<". Detailed description of the invention

[0036] According to the invention, the present invention relates to a process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis.

[0037] 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).

[0038] 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 to biogenic CO2 and to which a step of transforming CO2 into CO would have been associated. In addition, the hydrogen can also come from any existing process known to those skilled in the art. Examples 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 stages to eliminate impurities such as ammonia (NH3) or hydrogen sulfide (H2S).

[0039] 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 synthetic hydrocarbons obtained are mainly composed of paraffins but also contain olefins and oxygenates.

[0040] 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.

[0041] Preferably, the paraffinic effluent from the Fischer-Tropsch process is free from heteroatomic impurities such as, for example, sulfur, nitrogen or metals.

[0042] 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 this effluent or to a reconstituted cut of these two effluents.

[0043] The light fraction called cold condensate corresponds to the hydrocarbons in the gaseous state under the conditions of the Fischer-Tropsch reaction and the heavy fraction called waxes corresponds to the hydrocarbons in the liquid state under the conditions of the Fischer-Tropsch reaction.

[0044] The cold condensed light fraction 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.

[0045] 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 whose number of carbon atoms is greater than 8.

[0046] This heavy fraction has the particularity of being solid at room temperature due to its composition, formed essentially of n-paraffins having a high number of carbon atoms which can reach up to more than 150.

[0047] In the light fraction, the paraffin content is greater than 70% by weight relative to the total mass of said light fraction, 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.

[0048] Step a)

[0049] According to the invention, the process comprises a) step of hydrotreatment of at least a portion 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 velocity of between 0.1 and 40 h1.

[0050] 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 feedstock and preferably between 500 and 1500 normal liters of hydrogen per liter of feedstock, and at an hourly volumetric flow rate of between 0.25 and 20 h 1 and preferably between 0.5 and 10 h 1.

[0051] Said hydrotreatment step makes it possible to reduce the content of olefinic and unsaturated compounds as well as possibly to decompose the oxygenated compounds present in the effluent resulting from the Fischer-Tropsch synthesis.

[0052] 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 acquired from companies specializing in the manufacture and sale of catalysts.The shape of the supports and catalysts involved in step a) may be spherical, or extruded in the form of cylinders, trilobes or even quadrilobes according to any method known to those skilled in the art.

[0053] 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 least one Group VIII metal, preferably cobalt and nickel of the periodic table of elements is advantageously used. The concentration of non-noble Group VIII metal, when used, is advantageously from 0.01% to 15% by weight of equivalent relative to the finished catalyst and that of the Group VI metal is advantageously from 5% to 40% by weight of oxide equivalent relative to the finished catalyst. When a combination of Group VI and Group VIII metals is used, the catalyst is then preferably used in a reduced or sulfurized form.

[0054] 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 370°C into products having boiling points lower than 370°C is limited to 20% by weight, preferably is less than 10% by weight and even more preferably is less than 5% by weight.

[0055] 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 the fractionation step b) according to the invention.

[0056] Said step of eliminating at least part of the water can advantageously be carried out by any of the methods and techniques known to those skilled in the art, for example by drying, passing through a desiccant, flash or decantation.

[0057] Step b)

[0058] The process according to the invention comprises a step b) of fractionation in one or more steps 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 200°C, at least one middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C, and a heavy fraction having an initial boiling point of between 350 and 390°C.

[0059] Advantageously, step b) also makes it possible to separate said fraction with an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C into two fractions, the first having an initial boiling point of between 100 and 200°C and a final boiling point of between 250 and 285°C and the second having an initial boiling point of between 250 and 285°C and a final boiling point of between 350 and 390°C.

[0060] Said fractionation step can advantageously be carried out by atmospheric distillation and in certain cases by the combination of atmospheric distillation and vacuum distillation.

[0061] 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.

[0062] In a preferred embodiment, two middle distillate fractions are separated during fractionation step b), a light middle distillate fraction whose initial boiling point is between 100 and 200°C and whose final boiling point is between 200 and 250°C, and a heavy middle distillate fraction whose initial boiling point is between 200 and 250°C and whose final boiling point is between 350 and 390°C. In this embodiment, only the heavy middle distillate fraction is treated in said hydroisomerization step e). The light middle distillate fraction is advantageously not sent to said hydroisomerization step e) according to the invention.

[0063] The light middle distillate fraction from fractionation step b) can then advantageously be mixed with the middle distillate fraction from fractionation step h), the mixture then constituting the final middle distillate cut obtained by said process and having the required specifications.

[0064] 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).

[0065] Step c)

[0066] According to the invention, the process comprises a step c) of hydrocracking at least a portion 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°C and 450°C, at a pressure of between 0.2 and 15 MPa, at a space velocity of between 0.1 h 1 and 10 h *, 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 lower, at 0.4 and preferably less than 0.3, preferably less than 0.2, Tx being between 350 and 390°C, preferably between 360 and 380°C.

[0067] Preferably, step c) operates at a temperature of between 280 and 450°C, and even more preferably between 320 and 420°C, at a pressure of between 0.5 and 10 MPa, more preferably between 1 and 9 MPa, at a space velocity of between 0.2 and 7 h1, and more preferably between 0.5 and 5 h1, and with a hydrogen flow rate adjusted to obtain a ratio of between 150 and 1500 normal liters of hydrogen per liter of feedstock and more preferably between 300 and 1500 normal liters of hydrogen per liter of feedstock.

[0068] The heavy fraction contained in the feedstock of step c) advantageously has an initial boiling temperature of between 350 and 390°C, preferably between 360 and 380°C.

[0069] 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.

[0070] Preferably, said hydrocracking catalyst comprises at least one noble metal from group VIII chosen from platinum and palladium, taken alone or as a mixture, preferably active in their reduced form.

[0071] 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.

[0072] 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.

[0073] Advantageously, the acid support preferably from Bronsted comprises and preferably consists of silica alumina and / or zeolite.

[0074] In the case where the Bronsted acid support 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.

[0075] Preferably the Bronsted acid support comprises and preferably consists of silica-alumina.

[0076] Optionally, said support of the hydrocracking catalyst 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 (A12O3), clays, titanium oxide (TiO2), boron oxide (B2O 3) and zirconia (ZrO2) 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.

[0077] 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 support, without any other binder.

[0078] 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 even more preferably between 20% and 70% and between 22% and 45%. This silica content is perfectly measured using X-ray fluorescence.

[0079] Several preferred catalysts used in the hydrocracking step c) of the process according to the invention are described below.

[0080] 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:

[0081] - a BET specific surface area of ​​100 to 500 m2 / g, preferably between 200 and 450 m2 / g and very preferably between 200 and 300 m2 / g,

[0082] - an average diameter of the mesopores measured by mercury porosimetry included between 4 and 12 nm, preferably between 4 and 11 nm and very preferably between 5 and 11 nm,

[0083] - a total pore volume measured by mercury porosimetry 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,

[0084] - a volume of macropores, the diameter of which is greater than 50 nm, less than 0.02 ml / g.

[0085] - a content of alkaline or alkaline-earth compounds of less than 300 ppm mass and preferably less than 200 ppm mass.

[0086] 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. The pore volumes are also obtained by mercury porosimetry.

[0087] 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.

[0088] 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 distribution of the noble metal represents the distribution of the metal inside 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 clearly 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.

[0089] 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.

[0090] 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 middle distillate, and also to selectively produce middle distillate compared to naphtha, which makes it possible to maintain low conversion rates per pass.

[0091] According to the invention, the operating conditions of said step c) are adjusted so that the conversion by passes into products having boiling points greater than or equal to a temperature Tx into products with lower boiling points at this same temperature Tx is less than 0.5, preferably less than 0.4 and preferably less than 0.3, preferably less than 0.2, Tx being between 350 and 390°C, preferably between 360 and 380°C.

[0092] The temperature Tx, and the resulting conversion rate, is (are) preferably chosen so as to maximize the yield of middle distillates from the process according to the invention while respecting the specifications of the middle distillate.

[0093] The effect expected by the operating conditions applied and the conversion per pass that it generates is to limit the overcracking into light products such as gas or naphtha, the latter not being desired.

[0094] When the increase in temperature necessary to preserve the conversion per target pass becomes too great and the selectivity towards light products, gas and / or gasoline increases, the catalytic cycle can be interrupted and the catalyst is discharged then replaced by a new fresh catalyst with which the expected effect will be found again.

[0095] The conversion per pass is defined as:

[0096] Conversion of Tx+ into Tx = [(% mass of effluent Tx) - (% mass of load Tx)] / [100 - (% mass of load Tx)],

[0097] with: - % mass of Tx effluent: mass percentage of compounds having boiling points lower than Tx in the hydrocracking effluent,

[0098] - % mass of Tx charge: mass percentage of compounds having points boiling points lower than Tx in the hydrocracking feedstock.

[0099] 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 middle distillate 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 and preferably lower than 0.3, more preferably lower than 0.2.

[0100] Step d)

[0101] The process according to the invention comprises a step d) of fractionating at least all or part of the effluent from step c).

[0102] Preferably, all of the effluent from step c) is fractionated in step d).

[0103] In another embodiment, the effluent from step c) undergoes a purge before being fractionated in step d).

[0104] 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 boiling point between 100 and 200°C, at least one middle distillate fraction having an initial boiling point between 100 and 200°C and a final boiling point between 350 and 390°C, and a residual heavy fraction having an initial boiling point between 350 and 390°C, said fractionation steps b) and d) advantageously being able to be carried out in the same fractionation unit or in separate units.

[0105] Said fractionation step can advantageously be carried out by atmospheric distillation and in certain cases by the combination of atmospheric distillation and vacuum distillation.

[0106] 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.

[0107] Advantageously, step d) also makes it possible to separate said fraction with an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C into two fractions, the first having an initial boiling point of between 100 and 200°C and a final boiling point of between 250 and 285°C and the second having an initial boiling point of between 250 and 285°C and a final boiling point of between 350 and 390°C.

[0108] Said fractionation steps b) and d) can advantageously be carried out in the same fractionation unit or in separate units and preferably fractionation steps b) and d) are carried out in the same unit.

[0109] 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.

[0110] In a preferred embodiment in which fractionation steps b) and d) are carried out in the same unit, the effluent from step c) undergoes a purge before being fractionated. [YES] Step e)

[0112] According to the invention, the process comprises a step e) of hydroisomerization of at least a portion of the middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C from fractionation step b) and of all or part of the middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C from 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 between 200 and 450°C, a pressure between 1 and 15 MPa, a space velocity between 0.1 and 10 h1, a hydrogen flow rate adjusted to obtain a ratio between 100 and 2000 normal liters of hydrogen per liter of feed to produce a hydroisomerized effluent.

[0113] 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 h1, a hydrogen flow rate adjusted to obtain a ratio of between 150 and 2000 normal liters of hydrogen per liter of feed and preferably between 150 and 1500 normal liters of hydrogen per liter of feed.

[0114] According to a first variant, all of the middle distillate fractions obtained after fractionation(s) b) and d) are sent to the hydroisomerization step e).

[0115] According to a second variant, in the case where fractionation steps b) and d) are carried out in separate fractionation units, two middle distillate fractions are advantageously separated during fractionation step d), a light kerosene fraction having an initial boiling point of between 120 and 160°C and a final boiling point of less than or equal to 260°C, preferably less than 250°C and very preferably less than 230°C and a heavy middle distillate fraction having an initial boiling point of more than 230°C, and a final boiling point of between 350 and 390°C.

[0116] In this embodiment, only the heavy middle distillate fraction from fractionation step d) is sent to said hydroisomerization step e), mixed with the middle distillate 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).

[0117] According to a third variant, in the case where fractionation steps b) and d) are carried out in a single unit, two fractions are advantageously separated during the fractionation step, a light kerosene fraction having an initial boiling point of between 120 and 160°C and a final boiling point of less than or equal to 260°C, preferably less than 250°C and very preferably less than 230°C and a heavy middle distillate fraction having an initial boiling point of more than 230°C, and a final boiling point of between 350 and 390°C.

[0118] In this embodiment, only the heavy middle distillate fraction from fractionation step b) is sent to said hydroisomerization step e). The light kerosene fraction is advantageously not sent to said step e). hydroisomerization according to the invention and is mixed directly with the kerosene fraction from step h) of final fractionation.

[0119] In a fourth embodiment, in the case where fractionation steps b) and d) are carried out in separate fractionation units, a fraction produced during step d) having an initial boiling point of between 250 and 285°C and a final boiling point of between 350 and 390°C is directly extracted without being injected into hydroisomerization step e).

[0120] In a fifth embodiment, in the case where fractionation steps b) and d) are carried out in a single unit, a fraction produced during step b) having an initial boiling point of between 250 and 285°C and a final boiling point of between 350 and 390°C is directly extracted without being injected into hydroisomerization step e).

[0121] 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.

[0122] 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-211, taken alone or in a mixture, zeolites of structural type *BEA preferably chosen from Beta and Tschemichite) and zeolite IZM-2, taken alone or in a mixture.

[0123] The hydro / dehydrogenating function

[0124] Preferably the noble metal from group VIII of the catalyst used in step e) is chosen from palladium and platinum and is preferably platinum.

[0125] 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.

[0126] According to one or more embodiments, the content of noble metal from group VIII, and preferably the content of platinum, in the catalyst used in step c) is between 0.01% and 4% by weight, preferably between 0.05% and 2% by weight, relative to the total weight of said catalyst.

[0127] The catalyst used in step c) may also advantageously further comprise at least one additional metal chosen from the group formed by the metals of groups 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.

[0128] Preferably, the content of at least one additional metal in the catalyst used in step c) is between 0.01% and 2% by weight, preferably between 0.05% and 1% by weight, relative to the total weight of said catalyst.

[0129] According to one embodiment, the sulfur content in the hydroconversion catalyst is such that the ratio of the number of moles of sulfur to the number of moles of the at least one metal 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 hydroconversion catalyst. According to one or more embodiments, the presence of sulfur in the catalyst comes from potentially present impurities, such as for example in the alumina binder. According to another embodiment, the catalyst does not contain sulfur.

[0130] The hydroisomerizing function.

[0131] 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 mixture) and / or *BEA (e.g. chosen from Beta or Tschemichite). 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.

[0132] According to the invention, 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.

[0133] Preferably, the catalyst contains at least one zeolite chosen from IZM-2 and ZSM-12, taken alone or as a mixture.

[0134] 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.

[0135] 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.

[0136] Zeolite IZM-2 is a crystallized microporous 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 crystallized microporous solid whose crystal structure is described on the website of the International Zeolyst Association (http: / / www.iza-structure.org / ). It is a one-dimensional zeolite with 12 MR, its structural code is MTW. A 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, DOI10.1016 / 0144-2449(87)90015-7).

[0137] 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.

[0138] The binder.

[0139] 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.

[0140] 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, i.e., so as to ensure the addition to 100% by weight of the elements constituting the catalyst used in step e).

[0141] 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 molecular sieve with the binder or a precursor of the binder, such as for example boehmite, which by heat treatment is transformed into alumina.

[0142] A preferred 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.

[0143] Another preferred 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.

[0144] According to a preferred embodiment, the catalyst used in step e) more particularly comprises, and preferably consists of:

[0145] - from 1% to 90%, preferably from 3% to 80% and even more preferably from 4% 60% by weight of zeolite;

[0146] - from 0.01% to 4% by weight, preferably from 0.05% to 2% by weight of at least one metal of the group VIIIB, preferably platinum;

[0147] - optionally from 0.01% to 2% by weight, preferably from 0.05% to 1% by weight of at least minus one additional metal selected from the group formed by the metals of groups IIIA, IVA and VIIB; - 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 - optionally at least one binder, preferably alumina, ensuring the 100% complement in the catalyst, relative to the total weight of the catalyst from step e).

[0148] 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, beads, 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 beads in a rotating beader or drum, drop coagulation, "oil-drop", "oil-up", or any other known method of agglomeration of a powder containing alumina and possibly other ingredients chosen from those mentioned above.

[0149] 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.

[0150] The extrusion can advantageously be carried out by any conventional, commercially available tool. The paste resulting from the kneading is advantageously extruded through a die, for example using a piston or a single-screw or double-screw extrusion machine. This extrusion step can advantageously be carried out by any method known to those skilled in the art.

[0151] 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.

[0152] Preferably, the drying is carried out under a flow of air. Said drying can also be advantageously 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.

[0153] Said support, optionally dried, then preferably undergoes a calcination step.

[0154] 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, in a licked bed or in a static atmosphere. For example, the furnace used can be a rotating rotary furnace or be a vertical furnace 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.

[0155] Post-calcination treatments may optionally be carried out, so as to improve the properties of the support, for example textural.

[0156] 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 of between 0.1 and 25 MPa. For example, a reduction may comprise a two-hour stage at 150°C followed by a temperature rise 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 m3 hydrogen / m3 catalyst and the total pressure may be kept constant at 0.1 MPa. Any reduction method may advantageously be considered, either in situ (the reduction of the catalyst is carried out in the same unit where the catalytic reaction is 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).

[0157] The operating conditions implemented in step e) make it possible to obtain an isomerized effluent containing little or no olefinic compounds and oxygenated compounds.

[0158] The isomerized effluent from step e) is then sent to fractionation step h) to obtain at least one middle distillate fraction.

[0159] Step f)

[0160] According to the invention, the process comprises the recycling of all or part of the residual heavy fraction having an initial boiling point of between 350 and 390°C in the hydrocracking step c).

[0161] In the variant where steps b) and d) are carried out in a single unit, the hydrocracked effluent from step c), optionally purged, is advantageously recycled into said common fractionation unit in a mixture with all or part of the effluent from hydrotreatment step a).

[0162] 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.

[0163] 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.

[0164] 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 of between 350 and 390°C.

[0165] Preferably, all or part of said residual heavy fraction, optionally purged, is advantageously recycled directly in said hydrocracking step c).

[0166] The recycling of the residual heavy fraction is a key step of said invention because this recycling step makes it possible to convert at least a portion of the residual heavy fraction into at least one middle distillate fraction. This makes it possible to increase the middle distillate yield of the process. 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.

[0167] Said recycle 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.

[0168] The recycle rate is defined as the ratio of the mass flow rate entering hydrocracking step c), i.e. the cumulative flow rate of the heavy fraction from step b) only hydrotreated in step a) and the residual heavy fraction from step b) or step d) originating from hydrocracking step c) to the flow rate of the heavy fraction from step b) only hydrotreated in step a). It can be calculated directly from the conversion per pass previously defined according to the following formula:

[0169] Recycle Tx = 1 / (Conversion of Tx+ to Tx)

[0170] The recycle rate is then advantageously between 2 and 20, preferably between 3 and 10 and even more preferably between 4 and 6.

[0171] Step g)

[0172] 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).

[0173] Generally, this is a part of the residual heavy fraction and advantageously the heaviest fraction of said residual heavy fraction.

[0174] The implementation of the purge makes it possible to maximize the quantity of residual heavy fraction recycled to the hydrocracking step c) while maintaining the performance of the process, any heavy compounds refractory to hydrocracking likely to be accumulated during the recycle steps being able to be eliminated by the purge.

[0175] 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.

[0176] In another embodiment, the purge 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. Preferably, the purge corresponds to less than 10% by weight of the total feed entering the hydrotreatment step a), very preferably less than 7% by weight, very preferably less than 5% by weight and even more preferably less than 3% by weight.

[0177] The purge can be carried out by any methods 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 370°C, very preferably greater than 450°C and even more preferably greater than 540°C.

[0178] The effluent from step c), deduced purge is then fractionated.

[0179] Step h)

[0180] 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 100 and 200°C, and at least one middle distillate fraction having an initial boiling point of between 100 and 200°C.

[0181] Fractionation step h) advantageously makes it possible to separate the middle distillate cut into two cuts, a kerosene cut having an initial boiling point of between 100 and 200°C, preferably between 120 and 180°C, very preferably between 125 and 160°C and a final boiling point of between 220 and 320°C, preferably between 230 and 310°C, very preferably between 250 and 300°C and a diesel cut having an initial boiling point of between 230 and 300°C, preferably between 230 and 310°C, very preferably between 250 and 300°C and a final boiling point of between 350 and 390°C, preferably between 360 and 380°C.

[0182] At the end of this fractionation step h), the kerosene obtained has all the specifications required to be used as sustainable aviation fuel according to current standards. The diesel obtained also has all the specifications of a road fuel.

[0183] 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 middle distillate fraction comes from said fractionation unit, said light middle distillate fraction is mixed with the middle distillate fraction from final fractionation step h).

[0184] In the preferred embodiment where steps b) and d) are carried out in separate fractionation units, and in the case where one or more light middle distillate fraction(s) is (are) derived from said fractionation steps b) and / or d), said light middle distillate fraction(s) is (are) mixed with the middle distillate fraction resulting from the final fractionation step h).

[0185] The process according to the invention makes it possible to obtain a final middle distillate cut, consisting of the middle distillate fraction resulting from fractionation step h) only or optionally consisting of said middle distillate fraction resulting from fractionation step h) in a mixture with a light middle distillate fraction according to the preferred embodiments described above.

[0186] A mass balance makes it possible to demonstrate excellent selectivity of the process towards this final middle distillate cut, the unwanted cuts being produced in reduced proportions compared to 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 middle distillate has been reduced to a 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 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 and then on a hydroisomerization catalyst selective towards isomerization at the expense of cracking. Description of the figures.

[0187] [Fig.l]

[0188] [Fig.l] illustrates embodiment no. 1 in which steps b) and d) are carried out in the same fractionation unit and where all of the middle distillate resulting from the fractionation is treated in the hydroisomerization step e).

[0189] [Fig 1 bis]

[0190] [Fig.l] bis illustrates embodiment no. 1 in which steps b) and d) are carried out in the same fractionation unit and where the entire middle distillate 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.

[0191] [Fig.2]

[0192] [Fig.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 fraction of the middle distillate resulting from said fractionation is treated in the hydroisomerization step e), the so-called light fraction of the middle distillate also resulting from said fractionation is directly mixed with the kerosene from fractionation step h) without going through step e).

[0193] [Fig.3]

[0194] [Fig.3] illustrates embodiment No. 3 in which steps b) and d) are carried out in two separate fractionation units and where all of the middle distillate from step d) is treated in the hydroisomerization step e).

[0195] [Fig.4]

[0196] [Fig.4] illustrates embodiment No. 4 in which steps b) and d) are carried out in two separate fractionation units and where only the so-called heavy fraction of the middle distillate from step d) is treated in the hydroisomerization step e), the so-called light fraction of the middle distillate also from step d) is directly mixed with the kerosene from the fractionation step h) without going through step e).

[0197] [Fig.5]

[0198] [Fig.5] illustrates embodiment No. 5 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 said 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 fraction resulting from fractionation step h) without going through step e) and the diesel fraction is also isolated from said fractionation directly without going through step e).

[0199] [Fig.6]

[0200] [Fig.6] illustrates embodiment No. 6 in which steps b) and d) are carried out in two separate fractionation units and where only the so-called heavy kerosene fraction from step d) is treated in the hydroisomerization step e), the so-called light kerosene fraction also from step d) is directly mixed with the kerosene fraction from fractionation step h) without going through step e) and the diesel fraction is also isolated at the end of step d) without going through step e).

[0201] [Fig.7]

[0202] [Fig.7] 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 middle distillate 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 middle distillate cut from the other cuts and, on the other hand, to separate the residual heavy cut which is then recycled in step c). Therefore, the effluent from step c) is never treated in the hydroisomerization step e).

[0203] In [Fig.l], 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 gaseous 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 a fractionation unit (B / D) into a gas cut 5, a naphtha cut 6, a middle distillate fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which a partially converted 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 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 middle distillate hydrocarbon cut 7 produced by the fractionation unit (B / D) is injected into a hydroisomerization unit (E) whose hydrocarbon effluent 10 is then fractionated in a fractionation unit (H). At the end of the fractionation in the unit (H), four streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12, a diesel stream 15 and finally a last gas stream 13.

[0204] In Figure 1Bis, 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 gaseous 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 a fractionation unit (B / D) into a gas cut 5, a naphtha cut 6, a middle distillate fraction 7 and a heavy fraction 8. A possible purge is carried out by withdrawing a part of the heavy fraction 8, to eliminate the purge 14.The heavy fraction, purge deduced, 8', is sent to a hydrocracking unit (C) during which a partially converted effluent 80 is produced, the latter is then injected jointly with the cut 4 resulting from hydrotreatment in the fractionation unit (B / D), the hydrocarbon streams 4 and 80 forming the stream 480. The middle distillate hydrocarbon cut 7 produced by the fractionation unit (B / D) is injected into a hydroisomerization unit (E) whose effluent . hydrocarbon 10 is then fractionated in a fractionation unit (H). At the end of the fractionation in the unit (H), four streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12, a diesel stream 15 and finally a last gas stream 13.

[0205] In [Fig.2], the synthesis gas 1 is directed to a Fischer- Tropsch (A'). At the outlet of 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 gaseous 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 / D) into a gas cut 5, a naphtha cut 6, a light middle distillate fraction 71, a heavy middle distillate fraction 7 and a heavy fraction 8.The heavy fraction is sent to a hydrocracking unit (C) during which a partially converted 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 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 heavy middle distillate fraction 7 produced by the fractionation unit (B / D) is injected into a hydroisomerization unit (E) whose hydrocarbon effluent 10 is then fractionated in a fractionation unit (H). At the end of the fractionation in the unit (H), four streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12, a diesel stream 16 and finally a last gas stream 13.The kerosene stream 11 is finally mixed in a mixing unit (H') with the light middle distillate fraction 71, leading to the production of the kerosene stream 15.

[0206] In [Fig.3], the synthesis gas 1 is directed to a Fischer- Tropsch (A'). At the outlet of 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 gaseous 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 a fractionation unit (B) into a gas cut 5, a naphtha cut 6, a middle distillate fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which a partially converted 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, a middle distillate flow. 17 and a residual heavy fraction 9. The residual heavy fraction 9 is then injected together with the cut 8 from the fractionation unit (B) into the hydrocracking unit (C), the hydrocarbon streams 8 and 9 forming the stream 89. The middle distillate hydrocarbon cut 7 produced by the fractionation unit (B) is mixed with the cut 17 from the fractionation unit (D) to form the stream 710, the latter then being injected into a hydroisomerization unit (E) whose hydrocarbon effluent 10 is then fractionated in a fractionation unit (H). At the end of the fractionation in the unit (H), four streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12, a diesel stream 18 and finally a last gas stream 13.

[0207] In [Fig.4], 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 gaseous 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 middle distillate fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which a partially converted 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 flow of heavy middle distillates 71 and a flow of light middle distillates 72 and a residual heavy fraction 9. The residual heavy fraction 9 is then injected together 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 middle distillate hydrocarbon cuts 7 and 71 produced respectively by the fractionation units (B) and (D) form the stream 710 which is then injected into a hydroisomerization unit (E) whose hydrocarbon effluent 10 is then fractionated in a fractionation unit (H). At the end of the fractionation in the unit (H), four streams are produced: a stream corresponding to the kerosene cut 11, a naphtha stream 12, a diesel stream 16 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 72, leading to the production of the kerosene stream 15.

[0208] In [Fig.5], 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 gaseous 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 / D) into a gas cut 5, a naphtha cut 6, a light kerosene fraction 71, a heavy kerosene fraction 7, a diesel fraction 16 and a heavy fraction 8. The heavy fraction 8 is sent to a hydrocracking unit (C) during which a partially converted 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 stream 9, purge deducted, is then injected jointly with the cut 4 resulting from hydrotreatment in the fractionation unit (B), the hydrocarbon streams 4 and 9 forming the stream 49.The heavy kerosene fraction 7 produced by the fractionation unit (B / D) is injected into a hydroisomerization unit (E) whose hydrocarbon effluent 10 is then fractionated 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 71, leading to the production of the kerosene stream 15.

[0209] In [Fig.6], 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 gaseous 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 middle distillate fraction 7 and a heavy fraction 8.The heavy fraction 8 is sent to a hydrocracking unit (C) during which a partially converted 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 and a diesel flow 16 and a residual heavy fraction 9. The residual heavy fraction 9 is then injected together 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 middle distillate hydrocarbon cuts 7 and 71 produced respectively by the fractionation units (B) and (D) form the stream 710 which is then injected into a hydroisomerization unit (E) whose hydrocarbon effluent 10 is then fractionated in a fractionation unit (H). At the end of the fractionation in the unit (H), four streams are produced: a stream corresponding to . the kerosene cut 11, a naphtha stream 12, a diesel stream 17 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 72, leading to the production of the kerosene stream 15. The diesel stream 16 is mixed in a mixing unit (H”) with a stream 17 from the fractionation unit H to produce a diesel stream 19.

[0210] In [Fig.7], 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 gaseous 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 middle distillate 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 middle distillate 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 fractionated in a fractionation unit (H). Following the fractionation in the unit (H), four streams are produced: a stream corresponding to the kerosene cut 15, a gas stream 11, a naphtha stream 12, a diesel stream 13 and a residual heavy fraction 91. The residual heavy fraction 91 is recycled into the hydrocracking unit (C) mixed with the heavy fraction 8 to form the stream 918.

[0211] The examples illustrate the invention without limiting its scope. EXAMPLES

[0212] Example 1: preparation of hydrotreatment catalyst Cl (compliant).

[0213] The catalyst is an industrial catalyst based on nickel, molybdenum and phosphorus on alumina with molybdenum oxide MoO3 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.

[0214] Example 2: preparation of hydrocracking catalyst C2 (compliant).

[0215] The silica-alumina powder is prepared according to the synthesis protocol described in patent EPI 415 712A. The quantities of orthosilicic acid and aluminum hydrate are chosen so as to have a composition of 70% by weight of alumina A12O3 and 30% by weight of silica SiO2 in the final solid.

[0216] This mixture is rapidly homogenized in a commercial colloidal mill in the presence of nitric acid so that the nitric acid content of the suspension at the mill outlet, i.e. 8% relative to the mixed silica-alumina solid. Then the suspension is dried conventionally in an atomizer in a conventional manner from 300°C to 60°C. The powder thus prepared is shaped in a Z arm in the presence of 8% nitric acid relative to the anhydrous product. Extrusion is carried out by passing the paste through a die equipped with orifices of diameter 1.4 mm. The extrudates thus obtained are dried in an oven at 120°C then calcined under a flow of dry air at 550°C then calcined at 850°C in the presence of water vapor.

[0217] The characteristics of the support thus prepared are as follows:

[0218] - an average mesopore diameter measured by mercury porosimetry of 8.0 nm,

[0219] - a total pore volume measured by mercury porosimetry of 0.49 ml / g,

[0220] - a mesoporous volume measured by mercury porosimetry of 0.48 ml / g,

[0221] - a volume of macropores, the diameter of which is greater than 50 nm less than 0.02 ml / g,

[0222] - a BET surface area of ​​235 m2 / g.

[0223] 50 grams of silica-alumina extrudates are then subjected to a step dry impregnation 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.

[0224] Example 3: Preparation of a C3 hydroisomerization catalyst (compliant).

[0225] Synthesis of IZM-2 zeolite.

[0226] The IZM-2 zeolite was synthesized in accordance with the teaching of patent FR 2 918 050 B. A colloidal suspension of silica known under the commercial term Ludox HS-40 marketed by Aldrich, is incorporated into a solution composed of sodium hydroxide (Prolabo), structuring agent 1,6bis(methylpiperidinium)hexane dibromide, aluminum hydroxide (Aldrich) and deionized water. The molar composition of the mixture is as follows: 1 SiO2; 0.0060 Al2O3; 0.1666 Na2O; 0.1666 l,6bis(methylpiperidinium)hexane; 33.3333 H2O. The mixture is stirred vigorously for half an hour. The mixture is then transferred, after homogenization, into a PARR type autoclave. The autoclave is heated for 5 days at 170°C with stirring on a rotating spit (30 rpm). The product obtained is filtered, washed with deionized water to reach a neutral pH and then dried overnight at 100°C in an oven. The solid is then introduced into a muffle furnace to be calcined to remove the structuring agent.The calcination cycle includes a temperature rise to 200°C, a two-hour hold at this temperature, a temperature rise to 550°C followed by an eight-hour hold at this temperature and . finally a return to room temperature. The temperature increases 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 increase ramp of 2°C / min) in a crossed bed under dry air (2 normal liters per hour and per gram of solid). The solid thus obtained was analyzed by X-ray diffraction and identified as consisting of IZM-2 zeolite.Characterizations using 27Al NMR, X-ray fluorescence and ICP methods provide the following results for IZM-2:

[0227] - weight percentage of hexacoordinated aluminum atoms A1VI: 5%, .

[0228] - ratio of the number of moles of silicon divided by the number of moles of aluminum network, in mole / mole, Si / Al: 72,

[0229] - ratio of the number of moles of sodium divided by the number of moles of aluminum network, in mole / mole, Na / Al: 0.03.

[0230] Preparation of the IZM-2 / alumina support.

[0231] The IZM-2 / alumina support is obtained by kneading and extruding the IZM-2 zeolite with a Pural SB3 type alumina gel. The kneaded 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.

[0232] Impregnation of platinum on the IZM-2 / alumina support.

[0233] The 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 a flow of dry air (2 normal liters per hour and per gram of solid) in a tubular furnace under the following conditions:

[0234] - temperature rise from ambient to 450°C at 5°C / min;

[0235] - two-hour stage at 450°C;

[0236] - descent to ambient.

[0237] 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.

[0238] Example 4: Preparation of a C4 hydroisomerization catalyst (compliant).

[0239] Zeolite ZSM-12.

[0240] The ZSM-12 zeolite is supplied by the company Zeolyst. The solid was analyzed by X-ray diffraction and identified as consisting of ZSM-12 zeolite. Characterizations using 27Al NMR, X-ray fluorescence and ICP methods provide the following results for ZSM-12:

[0241] - weight percentage of hexacoordinated aluminum atoms A1VI: 0%;

[0242] - ratio of the number of moles of silicon divided by the number of moles of aluminum network, in mole / mole, Si / Al: 43;

[0243] - ratio of the number of moles of sodium divided by the number of moles of aluminum network, in mole / mole, Na / Al: 0.009.

[0244] Preparation of the ZSM-12 / alumina support.

[0245] The ZSM-12 / alumina support is obtained by kneading and extruding the ZSM-12 zeolite with a Pural SB3 type alumina gel. The kneaded 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 traversed 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.

[0246] Impregnation of platinum on the ZSM-12 / alumina support.

[0247] The 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 a flow of dry air (2 normal liters per hour and per gram of solid) in a tubular furnace under the following conditions:

[0248] - temperature rise from ambient to 450°C at 5°C / min;

[0249] - two-hour stage at 450°C;

[0250] - descent to ambient.

[0251] 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.

[0252] Example 5: Process for the production of kerosene and diesel not in accordance with the invention.

[0253] In this example illustrated in [Fig.7], a fractionation of the feed into at least one gaseous fraction, one naphtha fraction, one middle distillate fraction and one heavy fraction is carried out after the hydrotreatment step. Then the heavy fraction is treated in a hydrocracking step while the middle distillate fraction is treated in parallel in an isomerization step.

[0254] A paraffinic charge resulting from the Fischer-Tropsch synthesis has the characteristics provided in Table 1. This charge includes cold condensates and waxes.

[0255] [Table 1]

[0256] Table 1: Characteristics of the charge resulting from Fischer-Tropsch synthesis Simulated Distillation T (5% wt): 120 °CT (25% wt): 215 °CT (50% wt): 318 °CT (75% wt): 421 °CT (95% wt): 546 °C Compounds 370°C+ (by GC) 37% wt Density at 15°C 0.7679 Nitrogen content <2 ppm Sulfur content <2 ppm Total atomic oxygen content (by IR) 8700 ppm

[0257] The total atomic oxygen content in said fraction is measured by the infrared absorption technique described in patent application US2009 / 0018374A1.

[0258] The oxygen content represents the content of molecules comprising at least one oxygen atom present in the heavy fraction and is expressed as a weight percentage relative to the total pass of said fraction. The oxygen content is measured by gas chromatography.

[0259] Step a) hydrotreatment of the paraffinic feedstock.

[0260] The paraffinic feedstock is treated in a lost hydrogen through-bed, i.e. without recycling the hydrogen, on the hydrotreatment catalyst Cl under operating conditions which allow the elimination of olefinic and oxygenated compounds as well as traces of nitrogen.

[0261] The selected operating conditions are as follows:

[0262] hourly volumetric flow rate WH (charge volume / catalyst volume / hour) = 2h-l

[0263] total working pressure: 7 MPa

[0264] hydrogen / charge ratio: 700 normal liters / liter

[0265] temperature: 330°C

[0266] Before testing, the catalyst Cl undergoes a reduction step under the following operating conditions:

[0267] pure hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst,

[0268] rise from ambient temperature to 120°C: 10°C / min,

[0269] one-hour plateau at 120°C,

[0270] rise from 120°C to 450°C at 5°C / min,

[0271] two-hour stage at 450°C,

[0272] pressure: 7 MPa

[0273] The hydrotreated effluent then undergoes a flash separation and decantation step. At the end of said separation, the 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.

[0274] 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 370°C+ fraction into the 370°C- fraction is negligible (less than 5% by weight). The characteristics of the heavy fraction after hydrotreatment and separation are indicated in Table 2.

[0275] [Table 2]

[0276] Table 2: characteristics of the effluent after hydrotreatment and separation. Simulated Distillation T (5% wt): 125 °CT (25% wt): 173 °CT (50% wt): 317 °CT (75% wt): 420 °CT (95% wt): 536 °C compounds 370°C+ (by GC) 36% wt density at 15°C 0.7821 nitrogen content < 2 ppm sulfur content < 2 ppm oxygen content (by IR) <500 ppm

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

[0278] 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 middle distillate cut with an initial boiling point of 120°C and a final boiling point of 370°C, and a heavy fraction with an initial boiling point above 370°C.

[0279] Step e) hydroisomerization of the middle distillate fraction from step b).

[0280] The middle distillate fraction from step b) constitutes the hydroisomerization feed sent to the hydroisomerization catalyst C4 in accordance with the invention. The hydroisomerization feed is treated in a lost hydrogen through-bed, i.e. without recycling hydrogen.

[0281] Before testing, the C4 catalyst undergoes a reduction step under the following operating conditions:

[0282] - pure hydrogen flow rate: 1600 normal liters per hour and per liter of catalyst,

[0283] - increase in ambient temperature to 120°C: 10°C / min,

[0284] - one-hour stage at 120°C,

[0285] - rise from 120°C to 450°C at 5°C / min,

[0286] - two-hour stage at 450°C,

[0287] - pressure: 6.5 MPa.

[0288] After reduction, the catalytic test is carried out under the following conditions:

[0289] - total pressure of 6.5 MPa,

[0290] - hydrogen to charge ratio of 800 normal liters / liter,

[0291] - hourly volumetric speed (WH) equal to 2 h1.

[0292] The reactor temperature is adjusted to achieve a middle distillate yield of 97%.

[0293] The yield of middle distillates is equal to (% of 120-370°C effluents) / (% of 120-370°C feed) * 100

[0294] with:

[0295] % of 120-370°C effluents: mass percentage of compounds having points boiling point between 120 and 370°C in the effluents leaving the hydroisomerization unit;

[0296] % of 120-370°C load: mass percentage of compounds having points boiling point between 120 and 370°C in the feed at the inlet of the hydroisomerization unit;

[0297] Step c) hydrocracking of the heavy fraction from step b).

[0298] The heavy fraction from step b) mixed with the residual heavy fraction from step h) below constitutes the hydrocracking feed sent to the catalyst hydrocracking C2 according to the invention. The reduction conditions of catalyst C2 are identical to those of example 5.

[0299] After reduction, the catalytic test is carried out under the following conditions:

[0300] - total pressure of 6.5 MPa,

[0301] - hydrogen to charge ratio of 800 normal liters / liter,

[0302] - hourly volumetric speed (WH) equal to 1 h1.

[0303] 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 equal to:

[0304] C(370°C+) = [ (% of 370°C effluents) - (% of 370°C load) ] / [ 100 - (% of 370°C load)]

[0305] with

[0306] % of 370°C effluents = mass percentage of compounds having points boiling points below 370°C in the effluents,

[0307] and

[0308] % of 370°C - charge = mass percentage of compounds having points boiling points below 370°C in the hydrocracking feedstock.

[0309] The reaction temperature is adjusted to 335°C so as to obtain a conversion level of the 370°C+ fraction equal to 0.35. This corresponds to a recycle rate equal to 2.86. The conversion level of the 370°C+ fraction was adjusted so as to obtain a crystal disappearance temperature value for the kerosene obtained after fractionation step h) equal to -40°C measured according to ASTM D5972.

[0310] Step h) fractionation of the hydrocracking effluent.

[0311] The hydrocracking effluent obtained at the end of step c) is fractionated at the same time as the hydroisomerization effluent from 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 285°C, a diesel fraction with an initial boiling point of 285°C and a final boiling point of 370°C, and a residual heavy fraction with an initial boiling point greater than 370°C. Said residual heavy fraction is then recycled to the hydrocracking step c).

[0312] Material balances carried out over 24 hours make it possible to calculate the yields in different cuts:

[0313] - light gas yield: mass of gas Ci to C4 / mass of paraffinic charge * 100,

[0314] - yield in naptha cut: mass of liquid effluent of initial boiling point 35°C and final 120°C / mass of paraffinic charge * 100,

[0315] - kerosene cut yield: mass of liquid effluent with initial boiling point 120°C and final 285°C / mass of paraffinic charge * 100,

[0316] - diesel cut yield: mass of liquid effluent with initial boiling point 285°C and final 370°C / mass of paraffinic charge * 100,

[0317] 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 and of steps b) and h) of fractionation; the mass of paraffinic feedstock corresponds to the quantity of paraffinic feedstock 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 feedstock from the Fischer-Tropsch synthesis and of step h) of joint fractionation of the hydrocracking and hydroisomerization effluents,the mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 285°C corresponds to the quantity of liquid effluent with an initial boiling point of 120°C and a final boiling point of 285°C generated during 24 hours at the end of step h) of joint fractionation of the hydrocracking and hydroisomerization effluents and the mass of liquid effluent with an initial boiling point of 285°C and a final boiling point of 370°C corresponds to the quantity of liquid effluent with an initial boiling point of 285°C and a final boiling point of 370°C generated during 24 hours at the end of step h) of joint fractionation of the hydrocracking and hydroisomerization effluents. The crystal disappearance point of said kerosene cut is then measured according to standard ASTM D5972. The cloud point of said diesel cut is then measured according to standard ASTM D5773. ,

[0318] Table 3 reports the yields in different cuts as well as the properties of the kerosene and diesel cuts obtained.

[0319] Example 6: Process for the production of kerosene and diesel not in accordance with the invention.

[0320] The same paraffinic charge from the Fischer-Tropsch synthesis is used as in Example 5.

[0321] Example 6 is illustrated in [Fig. 1] in the case where the purge is zero.

[0322] Step a) hydrotreatment of the paraffinic feedstock.

[0323] Same as example 5.

[0324] Step b) fractionation of the hydrotreated and separated paraffinic feedstock.

[0325] Same as example 5.

[0326] Step c) hydrocracking of the heavy fraction from step b).

[0327] The heavy fraction from step b) mixed with the residual heavy fraction from step d) below constitutes the hydrocracking feed sent to the catalyst hydrocracking C2 according to the invention. The reduction conditions of catalyst C2 are identical to those of example 5.

[0328] After reduction, the catalytic test is carried out under the following conditions:

[0329] - total pressure of 6.5 MPa,

[0330] - hydrogen to charge ratio of 800 normal liters / liter,

[0331] - hourly volumetric speed (WH) equal to 1 h1.

[0332] The hydrocracking feedstock is treated in a lost hydrogen through-bed, i.e. without hydrogen recycling. The conversion of the 370°C+ fraction is taken as equal to:

[0333] C(370°C+) = [ (% of 370°C- effluents ) - (% of 370°C load) ] / [ 100 - (% of 370°C load)]

[0334] with

[0335] % of 370°C effluents = mass percentage of compounds having points boiling points below 370°C in the effluents,

[0336] and

[0337] % of 370°C charge = mass percentage of compounds having points boiling points below 370°C in the hydrocracking feedstock.

[0338] The reaction temperature is adjusted to 345°C so as to obtain a conversion level of the 370°C+ fraction equal to 0.65. This corresponds to a recycle rate equal to 1.54.

[0339] Step d) fractionation of the hydrocracking effluent.

[0340] The hydrocracking effluent obtained at the end of 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 and a final boiling point of 120°C, a middle distillate cut with an initial boiling point of 120°C and a final boiling point of 370°C, and a residual heavy fraction with an initial boiling point above 370°C. Said residual heavy fraction is then recycled to the hydrocracking step b). 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.

[0341] Step e) hydroisomerization of the middle distillate fraction from step b).

[0342] The middle distillate fraction from step b) constitutes the hydroisomerization feed sent to the hydroisomerization catalyst C4 in accordance with the invention. The reduction conditions of the catalyst C4 are identical to those of example 5. The hydroisomerization feed is treated in a lost hydrogen through-bed, i.e. without recycling hydrogen.

[0343] After reduction, the catalytic test is carried out under the following conditions:

[0344] - total pressure of 6.5 MPa,

[0345] - hydrogen to charge ratio of 800 normal liters / liter,

[0346] - hourly volumetric speed (WH) equal to 2 h1.

[0347] 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).

[0348] Step h) fractionation of the hydroisomerization effluent.

[0349] The hydroisomerization effluent obtained at the end of step e) is fractionated by atmospheric distillation. This fractionation is carried out in an atmospheric distillation unit different from step d).

[0350] Material balances carried out over 24 hours make it possible to calculate the yields in different cuts:

[0351] - light gas yield: mass of gas Ci to C4 / mass of paraffinic charge * 100

[0352] - yield in naptha cut: mass of liquid effluent of initial boiling point 35°C and final 120°C / mass of paraffinic charge * 100

[0353] - kerosene cut yield: mass of liquid effluent with initial boiling point 120°C and final 285°C / mass of paraffinic charge * 100

[0354] - diesel cut yield: mass of liquid effluent with initial boiling point 285°C and final 370°C / mass of paraffinic charge * 100

[0355] 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 resulting from the Fischer-Tropsch synthesis, of step b) and d) of joint fractionation of the hydrotreatment effluent and separation of the paraffin feedstock resulting from the Fischer-Tropsch synthesis and the hydrocracking effluent and of step h) of fractionation of the hydroisomerization effluent; the mass of paraffinic feedstock corresponds to the quantity of paraffinic feedstock resulting 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) 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, the mass of liquid effluent with an initial boiling point of 120°C and a final boiling point of 285°C corresponds to the quantity of liquid effluent with an initial boiling point of 120°C and a final boiling point of 285°C generated during 24 hours at the end of h) of fractionation of the hydroisomerization effluent and the mass of liquid effluent with an initial boiling point of 285°C and a final boiling point of 370°C corresponds to the quantity of liquid effluent with an initial boiling point of 285°C and, final temperature of 370°C generated for 24 hours at the end of h) fractionation of the hydroisomerization effluent. The disappearance point of the crystals of said kerosene cut is then measured according to standard ASTM D5972. The cloud point of said diesel cut is then measured according to standard ASTM D5773.

[0356] Table 3 reports the yields in different cuts as well as the properties of the kerosene and diesel cuts obtained.

[0357] Example 7: Process for the production of kerosene and diesel in accordance with the invention.

[0358] Step a) hydrotreatment of the paraffinic feedstock.

[0359] Same as example 5.

[0360] Step b) fractionation of the hydrotreated and separated paraffinic feedstock.

[0361] Same as example 5.

[0362] Step c) hydrocracking of the heavy fraction from step b).

[0363] Identical to example 6, except that the reaction temperature is adjusted to 330°C so as to obtain a conversion level of the 370°C+ fraction equal to 0.3- This corresponds to a recycle rate equal to 3.3.

[0364] Step d) fractionation of the hydrocracking effluent.

[0365] Same as example 6.

[0366] Step e) hydroisomerization of the middle distillate fraction from step b).

[0367] The middle distillate fraction from step b) constitutes the hydroisomerization feed sent to the hydroisomerization catalyst C4 in accordance with the invention. The reduction conditions of the catalyst C4 are identical to those of example 6. The hydroisomerization feed is treated in a lost hydrogen through-bed, i.e. without recycling hydrogen.

[0368] After reduction, the catalytic test is carried out under the following conditions:

[0369] - total pressure of 6.5 MPa,

[0370] - hydrogen to charge ratio of 800 normal liters / liter,

[0371] - hourly volumetric speed (WH) equal to 2 h1.

[0372] 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).

[0373] Step h) fractionation of the hydroisomerization effluent.

[0374] Same as example 5.

[0375] The yields in different cuts are calculated identically to example 5, the disappearance point of the crystals of the kerosene cut is also measured according to the ASTM D5972 standard. Similarly the finding point is measured according to the ASTM D5773 standard.

[0376] Table 3 reports the yields in different cuts as well as the properties of the kerosene cut obtained.

[0377] Example 8: Process for the production of kerosene and diesel in accordance with the invention.

[0378] Step a) hydrotreatment of the paraffinic feedstock.

[0379] Same as example 7.

[0380] Step b) fractionation of the hydrotreated and separated paraffinic feedstock.

[0381] Same as example 7.

[0382] Step c) hydrocracking of the heavy fraction from step b).

[0383] Same as example 7.

[0384] Step d) fractionation of the hydrocracking effluent.

[0385] Same as example 7.

[0386] Step e) hydroisomerization of the middle distillate fraction from step b).

[0387] The middle distillate fraction from step b) constitutes the feed hydroisomerization feedstock sent to the hydroisomerization catalyst C3 according to the invention. The reduction conditions of the catalyst C3 are identical to those of the catalyst C4 of Example 6. The hydroisomerization feedstock is treated in a lost hydrogen cross-bed, i.e. without recycling hydrogen.

[0388] After reduction, the catalytic test is carried out under the following conditions:

[0389] - total pressure of 6.5 MPa,

[0390] - hydrogen to charge ratio of 800 normal liters / liter,

[0391] - hourly volumetric speed (WH) equal to 2 h1.

[0392] 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).

[0393] Step h) fractionation of the hydroisomerization effluent.

[0394] Same as example 7.

[0395] Table 3 reports the yields in different cuts as well as the properties of the kerosene and diesel cuts obtained.

[0396] [Table 3]

[0397] Table 3: Yields in different cuts and properties of the middle distillate cuts obtained for the different examples of the invention. Example 5 not compliant Example 6 not compliant Example 7 compliant Example 8 compliant Yield of light gases Ci to C4 3.0 3.3 2.9 3.1 Naphtha cut yield 25.1 26.7 24.2 24.1 Kerosene cut yield 53.2 58.1 52.5 52.6 Diesel cut yield 19.0 12.3 20.6 20.4 Middle distillate yield 72.2 70.4 73.1 73.0 Kerosene cut properties Crystal disappearance point (°C) -40 -40 -40 -40 Gasoline cut properties Cloud point (°C) -15 -16 -18 -17

[0398] The examples provided make it possible to underline the importance of carrying out the process according to the invention. In particular, the comparison of Example 7 according to the invention with Example 5 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 5, the yields of kerosene, but also of diesel, are lower. Finally, the comparison of Example 7 according to the invention with comparative Example 6 underlines the importance of placing oneself at low conversion on the hydrocracking reactor of step c) since in the opposite case the yields of kerosene and diesel are also lower. Finally, Examples 7 and 8 make it possible to demonstrate that the process according to the invention using two types of hydroisomerization catalysts according to the invention makes it possible to maximize the yield of kerosene.

Claims

1. Claims Process for the production of middle distillates and preferably diesel from a paraffinic feedstock produced by Fischer-Tropsh synthesis comprising at least the following steps and preferably consisting of: a) the hydrotreatment of at least part 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 litres of hydrogen per litre of feedstock, and at an hourly volumetric flow rate of between 0.1 and 40 h1, 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 200°C, at least one middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C, and a heavy fraction having an initial boiling point of between 350 and 390°C. c) hydrocracking all of said heavy fraction having an initial boiling point of between 350 and 390°C from 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 h 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 preferably lower, at 0.2, Tx being between 350 and 390°C, preferably between 360 and 380°C, d) Fractionation in one or more stages of at least all or part of the effluent from stage c), 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 200°C, at least one middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C, and a residual heavy fraction having an initial boiling point of between 350 and 390°C, said fractionation stages 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 middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°C from fractionation step b) and of all or part of the middle distillate fraction having an initial boiling point of between 100 and 200°C and a final boiling point of between 350 and 390°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 h1, a hydrogen flow rate adjusted to obtain a ratio of between 100 and 2000 normal liters of hydrogen per liter of charge, 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 of all or part of the residual heavy fraction having an initial boiling point above 350°C in hydrocracking step (c), g) possibly purging part of the hydrocracked effluent from hydrocracking step c), h) fractionating all or part of the hydroisomerized effluent from hydroisomerization step e) 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 200°C, and at least one middle distillate fraction having an initial boiling point of between 100 and 200°C.

2. Method according to claim 1 in which step b) also makes it possible to separate said fraction with an initial boiling point between 100 and 200°C and a final boiling point between 350 and 390°C into two fractions, the first having an initial boiling point between 100 and 200°C and a final boiling point between 250 and 285°C and the second having an initial boiling point between 250 and 285°C and a final boiling point between 350 and 390°C.

3. Process according to one of claims 1 or 2 in which said fractionation step is carried out by atmospheric distillation or by the combination of atmospheric distillation and vacuum distillation or by stripping the effluent from step a) prior to atmospheric fractionation.

4. Process according to one of the preceding claims in which said hydrocracking catalyst used in step c) 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.

5. Process according to one of the preceding claims in which the acid support of the catalyst used in the hydrocracking step c) comprises and preferably consists of silica alumina and / or zeolite chosen from zeolites 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.

6. Process according to one of claims 1 to 5 in which a hydrocracking catalyst used in hydrocracking step c) comprises at least one noble metal, said noble metal being platinum and a silica-alumina as an acidic solid support, without any other binder.

7. Method according to one of the preceding claims in which step d) also makes it possible to separate said initial point fraction boiling point between 100 and 200°C and final boiling point between 350 and 390°C in two fractions, the first having an initial boiling point between 100 and 200°C and a final boiling point between 250 and 285°C and the second having an initial boiling point between 250 and 285°C and a final boiling point between 350 and 390°C.

8. Process according to one of claims 1 to 7 wherein in the case where fractionation steps b) and d) are carried out in separate fractionation units, two middle distillate fractions are advantageously separated during fractionation step d), a light kerosene fraction having an initial boiling point of between 120 and 160°C and a final boiling point of less than or equal to 260°C, preferably less than 250°C and very preferably less than 230°C and a heavy middle distillate fraction having an initial boiling point greater than 230°C, and a final boiling point of between 350 and 390°C and only the heavy middle distillate fraction from fractionation step d) is sent to said hydroisomerization step e), mixed with the middle distillate fraction from fractionation step b).

9. Process according to one of claims 1 to 7 wherein in the case where fractionation steps b) and d) are carried out in a single unit, two fractions are separated during the fractionation step, a light kerosene fraction having an initial boiling point of between 120 and 160°C and a final boiling point of less than or equal to 260°C, preferably less than 250°C and very preferably less than 230°C and a heavy middle distillate fraction having an initial boiling point greater than 230°C, and a final boiling point of between 350 and 390°C and only the heavy middle distillate fraction from fractionation step b) is sent to said hydroisomerization step e).

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