WO2010112691A1

WO2010112691A1 - Method for the production of middle distillates, comprising the hydroisomerisation and hydrocracking of a heavy fraction originating from a fischer-tropsch effluent using a resin

Info

Publication number: WO2010112691A1
Application number: PCT/FR2010/000252
Authority: WO
Inventors: Aurélie DANDEU; Vincent Coupard
Original assignee: IFP Energies Nouvelles; Eni S.P.A.
Priority date: 2009-04-03
Filing date: 2010-03-24
Publication date: 2010-10-07
Also published as: FR2944028B1; FR2944028A1; US20120048775A1

Abstract

The invention relates to a method for the production of middle distillates from a paraffin feed produced by means of Fischer-Tropsch synthesis and divided into two fractions, namely a light fraction, called cold condensate, and a heavy fraction, called waxes. The method comprises the following steps: (a) the fractionation of the cold condensate fraction into a gas fraction C4- and an intermediate fraction having an initial boiling point of between 15 and 40 °C and a final boiling point of between 300 and 450 °C; (b) the passage of the intermediate fraction over at least one ion exchange resin; (c) the elimination of at least part of the water from the effluent originating from step (b); (d) the decontamination of the heavy fraction, called waxes, by means of passage over a guard bed; (e) the recombination of the purified intermediate fraction and the effluent from step (d) in order to obtain a purified fraction C5+; (f) the hydrogenation of the unsaturated compounds of the purified fraction C5+; (g) the hydroisomerisation/hydrocracking of the effluent from step (f); (h) the separation and recycling in the hydroisomerisation/hydrocracking step of the unreacted hydrogen and the light gases; and (i) the distillation of the effluent from step (h).

Description

PROCESS FOR THE PRODUCTION OF MEDIUM DISTILLATES BY HYDROISOMERIZATION AND HYDROCRACKING OF A HEAVY FRACTION FROM A FISCHER-TROPSCH EFFLUENT

IMPLEMENTING A RESIN

In the Fischer-Tropsch process, the synthesis gas (CO + H2) is catalytically converted into oxygenates and substantially linear hydrocarbons in gaseous, liquid or solid form, these products constitute the charge of the process according to the invention. The paraffinic feedstock produced by Fischer-Tropsch synthesis used in the process according to the invention is produced from a synthesis gas in the Fischer Tropsch process, the synthesis gas (CO + H2) is advantageously produced in three ways. . In a preferred embodiment, the synthesis gas (CO + H2) is produced from natural gas according to the GTL route, also known in the English-German gas-to-liquid terminology. In another preferred embodiment, the synthesis gas (CO + H2) is produced from coal according to the CTL route, also known in the Anglo-Saxon coal-liquid terminology.

In another preferred embodiment, the synthesis gas (CO + H2) is produced from biomass according to the BTL pathway, also called according to the Anglo-Saxon biomass-to-liquid terminology.

However, these products, mainly made of normal paraffins, can not be used as such, in particular because of their cold-holding properties that are not very compatible with the usual uses of petroleum fractions. For example, the pour point of a linear hydrocarbon containing 20 carbon atoms per molecule (boiling point equal to about 340 ^° C., that is often included in the middle distillate cut) is + 37 ° C. about which makes its use impossible, the specification being -15 ° C for diesel. Thus, the hydrocarbons from the Fischer-Tropsch process comprising mainly n-paraffins must be converted into more valuable products such as, for example, gas oil, kerosene, which are obtained, for example, after catalytic hydrocracking / hydroisomerization reactions. On the other hand, they may have a significant content of unsaturated compounds of olefinic type and oxygenated products (such as alcohols, carboxylic acids, ketones, aldehydes and esters). These oxygenated and unsaturated compounds are more concentrated in the light fractions. Thus, in the C5 + fraction corresponding to the products boiling at an initial boiling point of between 15 ° C. and 40 ° C., these compounds represent between 10-20% by weight of olefinic type unsaturated compounds and between 5-10% by weight. of oxygenated compounds.

These products are generally free of heteroatomic impurities such as sulfur, nitrogen but may contain small amounts of Fe, Co, Zn, Ni or Mo from the dissolution of catalyst fines by the carboxylic acids. These metals can form complexes with oxygenates. These products practically do not contain little or no aromatics, naphthenes and more generally cycles especially in the case of cobalt catalysts.

The hydrogenation of the olefinic unsaturated compounds present in the Fischer-Tropsch hydrocarbons is a strongly exothermic reaction. Thus, under the severe hydrocracking / hydroisomerization operating conditions, the transformation of said unsaturated compounds may have a negative impact on the hydrocracking step such as a thermal runaway of the reaction, a large coking of the catalyst or gum formation by oligomerization. In order to protect the hydrocracking step, a hydrotreatment step is performed under less severe conditions than those for hydrocracking. However, the impurities of the feedstock, the oxygenated compounds and the metals (Fe ₁ Co, Zn, Ni, Mo) have a detrimental effect not only on the activity of the hydrotreatment and hydrocracking catalysts but also on the stability of the feedstock. hydrotreatment catalyst. Indeed, in the hydrotreatment reactor, the operating temperature conditions are such that the oxygenated compounds do not decompose but adsorb on the catalyst and form coke. In the hydrocracking section, the severe operating conditions cause the decomposition of the oxygenated compounds into water, carbon monoxide (CO) and carbon dioxide (CO2), which are inhibitors of the acid functions (water) and the hydrogenating function (CO). , CO ₂ ) of the hydrocracking catalyst and which then modifies the activity and selectivity of middle distillate. Therefore, the presence of oxygenated alcohols or acids present in the feeds requires an increase in the temperature in the hydrotreatment and hydrocracking stage to compensate for the decrease in activity and to maintain the conversion. In addition, the carboxylic acids can extract the active particles from the hydrotreatment and hydrocracking catalysts, thereby decreasing the life of said catalysts. Similarly, the metals complexed by the oxygenated compounds are decomposed on the active site of said hydrotreatment and hydroisomerization / hydrocracking catalysts in the presence of hydrogen and very selectively poison the active sites of said catalysts.

One of the objectives of the invention is therefore to reduce the total oxygen content of the feedstock and thus to limit the inhibitory effects of the oxygenated compounds and thus to limit the increase in temperature to compensate for the drop in activity and to maintain the conversion on the two stages of hydrotreatment and hydrocracking. The Applicant therefore implemented upstream of the hydrotreatment step and in order to increase the life of the hydrotreatment catalyst and the hydrocracking catalyst a step allowing, simultaneously or not, to transform on a resin exchange of ions, the alcohols and carboxylic acids constituting the oxygenated compounds in ester and of capturing the metals complexed by said oxygenated compounds.

This step is followed by a separation of water before the hydrotreatment step which makes it possible to reduce the total oxygen content and thus limit the inhibitory effects of the oxygenated compounds and thus limit the increase in temperature to compensate for the decrease. of activity and maintain the conversion on the two hydrotreatment and hydrocracking stages. The separation of the water also makes it possible to wash and capture CO and CO2 dissolved in the charge, which are inhibitors.

State of the art

The Shell patent application (EP-583,836) describes a process for the production of middle distillates from a filler obtained by Fischer-Tropsch synthesis. In this process, the feed resulting from the Fischer-Tropsch synthesis can be treated in its entirety, but preferably the C4- fraction is withdrawn from the feedstock so that only the C5 + fraction boiling at a temperature above 15 ° C. be introduced in the subsequent step. Said feedstock is subjected to a hydrotreatment to hydrogenate the olefins and alcohols, in the presence of a large excess of hydrogen, so that the conversion of products boiling above 370 ^° C. into products with a lower boiling point is less than 20%. The hydrotreated effluent consisting of high molecular weight paraffinic hydrocarbons is preferably separated from the hydrocarbon compounds having a low molecular weight and in particular the C4- fraction before the second hydroconversion stage. At least part of the remaining C5 + fraction is then subjected to a hydrocracking / hydroisomerization step with a conversion of products boiling above 370 ⁰ C into products with a boiling point of at least 40% by weight. Neither the presence of impurities in the feed nor the presence of removal steps of these impurities are mentioned in this application. The Shell patent application (EP-583,836) therefore does not deal with the problem of eliminating impurities present in the feedstock from the Fischer Tropsch process.

The patent applications SASOL (WO06 / 005084) and WO06 / 005085 deal with the removal of complexed metals by the oxygenated compounds present in a paraffinic filler resulting from the Fischer Tropsch process. These patents describe the decomposition after addition of water in a hydrothermal conversion zone of these compounds. This decomposition followed by physical treatment to remove the metals after decomposition. These requests require the addition of water to the system, the presence of three stages (reaction, separation of water, filtration) and does not affect the oxygenates present in the load. The present invention makes it possible to reduce the number of required steps, to eliminate the addition of water and to simultaneously carry out a conversion of oxygenated compounds present in the feedstock.

More specifically, the present invention relates to a method for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis and divided into two fractions, a light fraction, called cold condensate, and a heavy fraction, called waxes, comprising the steps of: a) fractioning said light fraction, referred to as cold condensate into two fractions, a C4-gas fraction boiling at a temperature below 15 ° C and an intermediate fraction having an initial boiling point of between 15 and 40 ⁰ C and a final boiling point between 300 and 450 ⁰ C, b) passing on at least one ion exchange resin of said intermediate fraction, at a temperature between 50 ⁰ C and 150 ⁰ C, at a total pressure between 0.7 and 2.5 MPa, at an hourly space velocity of between 0.2 and 2.5 h-1; and c) removal of at least a portion of the water from the effluent from the Etap eb), d) decontaminating by passing on a first guard bed containing at least a guard bed catalyst of said heavy fraction called waxes, e) recombination of the purified intermediate fraction from step c) and the effluent from step d) to obtain a purified C5 + fraction, f) hydrogenation of the olefinic unsaturated compounds of at least a portion of the purified C5 + fraction from step e) in the presence of hydrogen and a hydrogenation catalyst, g) hydroisomerization / hydrocracking in the presence of hydrogen and a hydroisomerization / hydrocracking catalyst of at least a portion of the effluent from step f) h) separation and recycling in the step g) hydroisomerization / hydrocracking unreacted hydrogen and light gases, i) distillation of the effluent from step h).

Detailed description of the invention

Throughout the rest of the description, we will detail the various steps of the method according to the invention with reference to Figures 1 and 2 which show preferred embodiments of the method according to the invention without limiting the scope. The effluent from the Fischer-Tropsch synthesis unit is, at the output of the Fischer-Tropsch synthesis unit divided into two fractions, a light fraction, called cold condensate (1), and a heavy fraction, called waxes. (11).

The two fractions thus defined comprise water, carbon dioxide (CO ₂ ), carbon monoxide (CO) and unreacted hydrogen (H ₂ ). In addition, the light fraction, cold condensate, contains light hydrocarbon compounds C1 to C4, called C4- fraction, in the form of gas.

Step a)

According to step a) of the process according to the invention, the light fraction, called cold condensate is fractionated into two fractions, a gaseous fraction C4- boiling at a temperature below 15 ° C and an intermediate fraction having a point d initial boiling between 15 and 40 ⁰ C and a final boiling point between 300 and 450 ⁰ C. The light fraction, called cold condensate (1), enters a fractionation means (2). The fractionation means (2) may for example consist of methods well known to those skilled in the art such as a flash or flash tank, distillation or stripping. Advantageously, the fractionation means is a distillation column allowing the elimination, at the top of the column, of light and gaseous hydrocarbon compounds C1 to C4, called gas fraction C4- (3), corresponding to the products boiling at a temperature below 15. ⁰ C, preferably less than 1O ⁰ C and very preferably below 0 ⁰ C and to retrieve an intermediate fraction (5) corresponding to the C5-C30 hydrocarbons, having an initial boiling point ranging between 15 and 40 ⁰ C and preferably between 15 and 25 ° C and a final boiling point between 300 and 450 ⁰ C and preferably between 380 and 400 ⁰ C.

Said intermediate fraction (5) and the heavy fraction called waxes are then separately treated before being recombined so as to obtain a purified C5 + liquid fraction in line (15), corresponding to products boiling at an initial boiling point between 15 and 40 ⁰ C and preferably having a boiling point greater than or equal to 20 ⁰ C.

Preferably, prior to step b) of the process according to the invention, the intermediate fraction (5) having an initial boiling point of between 15 and 40 ° C and a final boiling point between 300 and 450 ⁰ C undergoes an optional decontamination step in a reactor (34), shown in Figure 2, by passage over an optional guard bed containing at least a guard bed catalyst. The embodiments of said optional decontamination step as well as the guard bed catalysts used are the same as those used in step d) of decontamination of the process according to the invention and are described below, in step d ).

Step b)

According to step b) of the process according to the invention, said intermediate fraction (5) having an initial boiling point of between 15 and 40 ^c C and a final boiling point between 300 and 45O ⁰ C, optionally previously decontaminated by passage over a guard bed, passes over at least one ion exchange resin for the esterification of alcohols and carboxylic acids to ester and / or the capture of dissolved metals in the feed, at a temperature between 50 ° ^C. C and 150 ^° C. and preferably between 80 ^° C. and 150 ^° C., at a total pressure of between 0.7 and 2.5 MPa, at an hourly space velocity of between 0.2 and 2.5 h -1,

Step b) according to the invention can advantageously be implemented according to two distinct embodiments, namely, either in a single reactor (6) on a single ion exchange resin advantageously used to simultaneously perform the esterification of the alcohols and carboxylic acids to ester and the uptake of dissolved metals in the feed, either in two different reactors (not shown in the figures) on two different ion exchange resins, one whose specific function is the esterification of alcohols and carboxylic acids and the other the uptake of dissolved metals in the feed.

According to a first embodiment, step b) advantageously consists in the passage of said intermediate fraction (5) on a single ion exchange resin in a single reactor (6) making it possible simultaneously to carry out the esterification of the alcohols and acids carboxylic esters and the uptake of dissolved metals into the feedstock.

Preferably, said resin is used at a temperature of between 80 ^° C. and 150 ° C. and preferably between 80 ° and 130 ^° C., at a pressure of between 1 and 2 MPa and preferably between 1 and 1.5 ° C. MPa and at an hourly volume velocity of between 0.5 and 2 h -1 and preferably between 0.5 and 1.5 h -1.

In this case, the oxygenated compounds, carboxylic acids and alcohols are adsorbed on the active sites of said resin and are esterified and the cationic and metallic compounds present in said C5 + liquid paraffinic fraction are removed by adsorption or by ion exchange. said resin, which makes it possible simultaneously to carry out the esterification of the alcohols and carboxylic acids to the ester and to capture the dissolved metals in the feedstock, advantageously comprises acidic sulphonic groups and is prepared by polymerization or co-polymerisation of vinyl aromatic groups followed by sulphonation said aromatic vinyl groups being selected from styrene, vinyl toluene, vinyl naphthalene, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinyl xylene, said resin having a degree of crosslinking of between 20 and 35%, preferably between 25 and 35% and preferably equal to 30% and an acidic strength determined by potentiometry during neutralization with a KOH ₁ solution of 0.2 to 6 mmol H + equivalent / g and preferably between 0, 2 and 2.5 mmol H + equivalent / g.

Said acid ion exchange resin advantageously contains between 1 and 2 terminal sulphonic groups per aromatic group. Preferably, said resin has a size between 0.15 and 1.5 mm. The size of the resin is the diameter of the sphere encompassing the resin particle. Resin size classes are measured by sieving on sieves adapted according to a technique known to those skilled in the art.

A preferred resin is a resin consisting of aromatic vinyl mono vinyl aromatic poly vinyl copolymers, and very preferably, copolymer of vinylbenzene and polystyrene having a degree of crosslinking of between 20 and 35% and preferably between 25 and 35% and preferably equal to 30% and an acidic force, representing the number of active sites of said resin, assayed by potentiometry during neutralization with a KOH solution, of between 0.2 and 6 mmol. H + equivalent / g and preferably between 0.2 and 2.5 mmol H + equivalent / g.

Another preferred resin which simultaneously allows the esterification of alcohols and carboxylic acids and the uptake of metals is a resin consisting of a polysiloxane grafted with alkylsulphonic acid groups (of the type -CH 2 -CH 2 -CH 2 -SO 3 H), which is of considerable size. between 0.5 and 1.2 mm and acid strength representing the number of active sites of said resin and assayed by potentiometry during neutralization with a KOH solution, is 0.4 to 1.5 mmol H + equivalent / boy Wut.

During this step and under these conditions, 95% of the carboxylic acids are esterified. The analysis of the acid conversion is given by difference in potash titration between the feedstock and the effluent according to a technique known to those skilled in the art. We for example, ASTM methods D 664, D 3242 or D 974 as a method for carrying out this analysis.

This resin may advantageously be used in a fixed bed between grids placed in an upward or downward tubular reactor. Preferably, said resin is used in an upflow reactor, the liquid being injected down the reactor at a surface speed sufficient to cause expansion of the resin bed without transporting or fluidizing it. This embodiment, with respect to the fixed bed, makes it possible to attenuate the effects of the clogging materials and to substantially increase the life cycle of the resin.

According to a second embodiment, step b) advantageously consists in the passage of said intermediate fraction (5), in two different reactors, on two different ion exchange resins, of a different nature, one having the specific function the esterification of alcohols and carboxylic acids and the other the uptake of dissolved metals in the feedstock.

Preferably, the reactor containing the ion exchange resin for capturing the metals is used upstream of the reactor containing the ion exchange resin for the esterification of alcohols and carboxylic acids.

In this case, the cationic and metallic compounds present in said fraction said intermediate fraction (5) are removed by adsorption or by ion exchange on a first ion exchange resin. This first resin specific for the capture of metals advantageously comprises acidic sulphonic groups and is advantageously prepared by polymerization or co-polymerization of vinyl aromatic groups followed by a sulphonation, said aromatic vinyl groups being advantageously chosen from styrene, vinyl toluene, naphthalene vinyl, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene, and vinyl xylene, said resin having a degree of crosslinking of between 1 and 20% and preferably between 2 and 8% and an acidic strength, representing the number of active sites of said resin, assayed by potentiometry during neutralization with a KOH solution, of between 1 and 15 mmol H + equivalent / g and preferably between 2.5 and 10 mmol H + equivalent / g.

Said first acidic ion exchange resin advantageously contains between 1 and 2 terminal sulphonic groups per aromatic group. Preferably, said resin has a size between 0.15 and 1.5 mm. The size of the resin is the diameter of the sphere encompassing the resin particle. Resin size classes are measured by sieving on sieves adapted according to a technique known to those skilled in the art.

Preferably, the first resin is a resin consisting of aromatic vinyl mono vinyl aromatic polymers and, very preferably, di-vinylbenzene copolymer and polystyrene having a degree of crosslinking of between 1 and 20% and an acidic force representing the number of active sites of said resin and assayed by potentiometry during the neutralization with a KOH solution, and 1 and 15 mmol H + equivalent / g and preferably between 2.5 and 10 mmol H + equivalent / boy Wut.

Preferably, said first resin is used at a temperature of between 5O ⁰ C and 110 ⁰ C and preferably between 80 and 110 ⁰ C, at a pressure of between 1 and 2 MPa and preferably between 1 and 1 , 5 MPa and at an hourly space velocity of between 0.2 and 1.5 h -1 and preferably between 0.5 and 1.5 h -1.

The effluent from the reactor containing said first resin specific for the capture of metals is then advantageously introduced into a second reactor located downstream of the first reactor and containing a second resin of different nature and specific to the esterification of the alcohols and carboxylic acids contained in said effluent.

The oxygenated compounds, carboxylic acids and alcohols adsorb to the active sites of said second resin and are esterified and the cationic and metal compounds present in said intermediate fraction (5) are removed by adsorption or ion exchange. Said second resin, which makes it possible to carry out the esterification of the alcohols and carboxylic acids to the ester, advantageously comprises acidic sulphonic groups and is advantageously prepared by polymerization or co-polymerization of vinyl aromatic groups followed by sulphonation. The vinyl aromatic groups are advantageously chosen from styrene, vinyl toluene, vinyl naphthalene, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinyl xylene, said second resin having a degree of crosslinking, ie a copolymer mass / polymer mass ratio advantageously between 20 and 35% and preferably between 25 and 35% and preferably equal to 30% and an acidic force, representing the number of active sites of said resin, assayed by potentiometry during the neutralization by a KOH solution, between 0.2 to 6 mmol H + equivalent / g and preferably between 0.2 and 6 mmol H + equivalent / g.

Said second acid ion exchange resin advantageously contains between 1 and 2 terminal sulphonic groups per aromatic group. Preferably, said second resin has a size of between 0.15 and 1.5 mm.

A second preferred resin is a resin consisting of aromatic vinyl mono vinyl and aromatic poly vinyl co-polymers, and very preferably, copolymer of di-vinyl benzene and polystyrene having a degree of crosslinking of between 20 and 35% and preferably between 25 and 35% and preferably equal to 30% and an acidic force, representing the number of active sites of said resin, assayed by potentiometry during the neutralization with a KOH solution, of between 0.2 to 6%. mmol H + equivalent / g and preferably between 0.2 and 6 mmol H + equivalent / g.

Another second preferred resin which simultaneously allows the esterification of alcohols and carboxylic acids and the uptake of metals is a resin consisting of a polysiloxane grafted with alkylsulphonic acid groups (of the -CH2-CH2-CH2-SO3H type), size between 0.5 and 1, 2 mm and acid strength representing the number of active sites of said resin and assayed by potentiometry during the neutralization with a KOH solution, is 0.4 to 1.5 mmol H + equivalent / boy Wut.

Preferably, said second resin is used at a temperature of between 80 ^° C. and 150 ^° C. and preferably between 80 ° and 130 ^° C., at a pressure of between 1 and 2 MPa and preferably between 1 and 1, 5 MPa and at an hourly space velocity of between 0.5 and 2 h -1 and preferably between 0.5 and 1.5 h -1.

During this step and under these conditions, 95% of the carboxylic acids are esterified. The analysis of the acid conversion is given by difference in potash titration between the feedstock and the effluent. For example, ASTM methods D 664, D 3242 or D 974 can be used as a method for carrying out this analysis.

These resins may advantageously be used in a fixed bed between grids placed in a tubular reactor ascending or descending. Preferably, said resins are used in an upflow reactor, the liquid being injected reactor bottom at a surface speed sufficient to cause expansion of the resin bed without transporting or fluidizing it. This embodiment, with respect to the fixed bed, makes it possible to attenuate the effects of the comatant materials and to increase substantially the life cycle duration of the resin.

In the case where said intermediate fraction (5) passes in two different reactors, on two different ion exchange resins, of different nature, one mainly producing the metal uptake, the other mainly performing the esterification, it is used preferably intermediate heating between the two steps. Preferably, the water formed during the esterification step of the acids and alcohols is removed from the first resin essentially making the uptake of metals to push the esterification reaction onto the second resin. The simultaneous addition of the intermediate reheating and the separation of the water generally favors the conversion of the carboxylic acids present in the charge.

The esterification reaction of the organic acids with the alcohols present in said intermediate fraction (5) produces water which is an inhibitor compound for hydrotreatment and hydrocracking catalysts requiring an increase in the severity of the operating conditions.

Step c)

The effluent from step b) then undergoes according to the invention a step of removing at least a portion of the water formed during said step b) and preferably all of the water formed, in a separator (8).

This water is acidic in nature because it advantageously contains the protons exchanged during the uptake of metals by the specific cation exchange resin placed upstream or by the single resin simultaneously allowing the esterification and the uptake of the metals.

This water may also contain CO and dissolved CO2 from the synthesis

Fischer Tropsch. The water is removed by the pipe (9).

It is also advantageously added in said step c) nitrogen (N2) or hydrogen (H2) type gas to remove more CO and CO2 dissolved by stripping.

In the case of the addition of hydrogen, it advantageously serves as makeup gas for the hydrotreating step.

This step also makes it possible to eliminate light ether-type products formed during the reaction of the alcohols on themselves. The removal of water can be carried out by all the methods and techniques known to those skilled in the art, for example by drying, passage on a desiccant, flash, decantation ....

The effluent from step c) of removing the water is then recombined, according to step e) of the process according to the invention, with the effluent from step d) described below.

Step d)

According to the invention, the heavy fraction called wax undergoes a decontamination step in a reactor (12) by passing on a guard bed containing at least a guard bed catalyst.

Indeed, the treated heavy fraction may optionally contain solid particles such as inorganic solids. It may optionally contain metals contained in hydrocarbon structures such as organometallic compounds that are more or less soluble. By the term fines is meant fines resulting from a physical or chemical attrition of the catalyst. They can be micron or sub-micron. These mineral particles then contain the active components of these catalysts without the following list being limiting: alumina, silica, titanium, zirconia, cobalt oxide, iron oxide, tungsten, rhuthenium oxide, etc. These solid minerals may be present under the calcined mixed oxide form: for example, alumina-cobalt, alumina-iron, alumina-silica, alumina-zirconia, alumina-titanium, alumina-silica-cobalt, alumina-zirconia-cobalt

Said heavy fraction may also contain metals within hydrocarbon structures, which may optionally contain oxygen or more or less soluble organometallic compounds. More particularly, these compounds may be based on silicon. It may be for example anti-foaming agents used in the synthesis process. For example, the solutions of a silicon compound of silicon type or silicone oil emulsion are more particularly contained in the heavy fraction. On the other hand, the catalyst fines described above may have a higher silica content than the catalyst formulation resulting from the intimate interaction between the catalyst fines and anti-foaming agents described above.

The problem that is then posed is to reduce the content of solid mineral particles and possibly reduce the content of harmful metal compounds for the hydroisomerization-hydrocracking catalyst.

Characteristics of the catalysts used in the guard beds The guard beds advantageously contain at least one catalyst.

The guard bed catalysts used in steps c) of the process according to the invention and in the optional step of decontaminating said intermediate fraction may be different and preferably identical. The said guard bed catalysts are described below.

Form of catalysts

The guard bed catalysts used in steps c) of the process according to the invention and in the optional step of decontaminating said intermediate fraction may advantageously be in the form of spheres or extrudates. It is however advantageous that the catalyst is in the form of extrudates with a diameter of between 0.5 and 5 mm and more particularly between 0.7 and 2.5 mm. The shapes are cylindrical (which can be hollow or not), cylindrical twisted, multilobed (2, 3, 4 or 5 lobes for example), rings. The cylindrical shape is preferably used, but any other shape may be used.

In order to remedy the presence of contaminants and / or poisons in the feed, the guard catalysts may, in another preferred embodiment, have more particular geometric shapes in order to increase their void fraction. The vacuum fraction of these catalysts is advantageously between 0.2 and 0.75. Their outer diameter may advantageously vary between 1 and 35 mm. Among the particular forms possible without this list being exhaustive: hollow cylinders, hollow rings, Raschig rings, serrated hollow cylinders, crenellated hollow cylinders, pentaring carts, multi-hole cylinders ...

Active Phase The said guard bed catalysts used in steps c) of the process according to the invention and in the optional step of decontaminating said intermediate fraction may advantageously have been impregnated with an active phase or not. Preferably, the catalysts are impregnated with a hydro-dehydrogenation phase. Very preferably, the CoMo or NiMo phase is used. Even more preferably, the NiMo phase is used.

Preferably, the supports of said guard bed catalysts are porous refractory oxides, preferably chosen from alumina and silica-alumina. The said guard bed catalysts may advantageously have macroporosity.

Said catalysts advantageously comprise a macroporous mercury volume for a mean diameter at 50 nm which is greater than 0.1 cm 3 / g, preferably of between 0.125 and 0.175 cm 3 / g and a total volume of greater than 0.60 cm 3 / g, of preferably between 0.625 and 0.8 cm3 / g, and is advantageously impregnated with an active phase, preferably based on nickel and molybdenum such as for example IΑCT961. In this preferred embodiment, the Ni content by weight of oxide is generally between 1 and

10% and the Mo content by weight of oxide is between 5 and 15%. The surfaces expressed in SBET of the supports of said catalysts vary between 30 m ² / g and 220 m ² / g.

In a first embodiment, the guard bed advantageously also comprises at least one other catalyst having a mercury volume for a pore diameter greater than 1 micron greater than 0.2 cm 3 / g and preferably greater than 0.5 cm 3 / and a mercury volume for a pore diameter greater than 10 microns greater than 0.25 cm 3 / g and preferably less than 0.4 cm 3 / g, said catalyst being advantageously placed upstream of the first guard bed catalyst described herein. above.

In a second embodiment, the guard bed advantageously also comprises at least one other catalyst having a mercury volume for a pore diameter greater than 50 nm, greater than 0.25 cm3 / g, the mercury volume for a pore diameter. greater than 100 nm, greater than 0.15 cm3 / g and a total pore volume greater than 0.80 cm3 / g.

Said guard bed catalyst and the catalyst according to the first embodiment can advantageously be combined in a mixed bed or a combined bed. Generally the impregnated active phase catalyst constitutes the majority of the guard bed and the catalyst according to the first preferred embodiment is added in addition to 0 to 50% by volume relative to the first catalyst, preferably from 0 to 30%, even more preferably from 1 to 20%. The combination of said guard bed catalyst and the catalyst according to the first embodiment does not limit the scope of the invention. In fact, catalysts that can be used in the guard beds can advantageously be used alone or in mixtures and chosen in a non-exhaustive manner from the catalysts marketed by Norton-Saint-Gobain, for example the MacroTrap® guard beds or the catalysts marketed by Axens. in the ACT family: ACT077, ACT935, ACT961 or HMC841, HMC845, HMC941, HMC945 or ACT645.

The preferred guard beds according to the invention are HMC and IΑCT961. It may be particularly advantageous to superpose these catalysts in at least two different beds of variable height. The catalysts having the highest void content are preferably used in the first catalytic bed or first catalytic reactor inlet. It may also be advantageous to use at least two different reactors for these catalysts.

Advantageously, a combination of said bed bed catalyst with the catalysts according to the first and second embodiments is also possible in a mixed bed or a combined bed. In this case, the catalysts are placed by decreasing rate of vacuum in the direction of flow.

After passing over said guard bed, the content of solid particles is less than 20 ppm, preferably less than 10 ppm and even more preferably less than 5 ppm. The soluble silicon content is less than 5 ppm, preferably less than 2 ppm and even more preferably less than 1 ppm.

Preferably, the effluent from step d) of decontamination by passing over a guard bed of said heavy fraction called wax passes, optionally, before the recombination step e) of the purified intermediate fraction resulting from the step c) (line 10) and said effluent from step d) over at least one ion exchange resin at a temperature between 5O ⁰ C and 150 ⁰ C and preferably between 80 and 150 ⁰ C at a total pressure between 0.7 and 2.5 MPa, at an hourly space velocity of between 0.2 and 2.5 h -1. This optional decontamination step may advantageously be carried out according to two distinct embodiments, namely, either in a single reactor (35) (shown in FIG. 2) on a single ion exchange resin advantageously used to simultaneously produce the esterification of the alcohols and carboxylic acids to the ester and the uptake of the dissolved metals in the feedstock, either in two different reactors (not shown in the figures) on two ion exchange resins of different nature, one having the specific function l esterification of alcohols and carboxylic acids and the other the uptake of dissolved metals in the feedstock.

The embodiments above and the resins used are the same as those used in step b) of the process according to the invention and are described above. Step e)

According to the invention, the purified intermediate fraction resulting from stage c) (line 10) and the effluent from stage d) of decontamination, optionally purified by passage over at least one ion exchange resin, (line 13) are recombined in the line (15) to obtain a purified C5 + fraction which constitutes the charge of step f) of hydrogenation of the olefinic type unsaturated compounds.

Step f)

Stage f) of the process according to the invention is a stage of hydrogenation of the olefinic type unsaturated compounds of at least a part and preferably all of the effluent resulting from stage e) of the process according to the invention. in the presence of hydrogen and a hydrogenation catalyst.

The effluent from step e) (line 15) of the process according to the invention is allowed in the presence of hydrogen (line 14) in a hydrogenation zone (16) containing a hydrogenation catalyst which has the objective of saturating the olefinic type unsaturated compounds present in said effluent.

Preferably, the catalyst used in step f) according to the invention is a non-cracking or slightly cracking hydrogenation catalyst comprising at least one metal of group VIII of the periodic table of the elements and comprising at least one support based on of refractory oxide.

Preferably, said catalyst comprises at least one Group VIII metal chosen from nickel, cobalt, ruthenium, indium, palladium and platinum and comprising at least one refractory oxide-based support chosen from alumina. and silica alumina.

Preferably, the group VIII metal is chosen from nickel, palladium and platinum and very preferably from palladium and platinum.

According to a preferred embodiment of step f) of the process according to the invention, the group VIII metal is chosen from palladium and / or platinum and the content of this metal is advantageously between 0.1% and 5%. % by weight, and preferably between 0.2% and 0.6% by weight relative to the total weight of the catalyst.

According to a very preferred embodiment of step f) of the process according to the invention, the Group VIII metal is palladium. According to another preferred embodiment of step f) of the process according to the invention, the metal of group VIII is nickel and the content of this metal is advantageously between 5% and 25% by weight, preferably between 7%. and 20% by weight based on the total weight of the catalyst.

The catalyst support used in step f) of the process according to the invention is a refractory oxide-based support, preferably chosen from alumina and silica-alumina and preferably alumina.

When the support is an alumina, it has a BET specific surface to limit the polymerization reactions on the surface of the hydrogenation catalyst, said surface being between 5 and 140 m ² / g.

When the support is a silica-alumina, the support contains a percentage of silica of between 5 and 95% by weight, preferably between 10 and 80%, more preferably between 20 and 60% and very preferably between 30 and 50%. a BET specific surface area of between 100 and 550 m ² / g, preferably between 150 and 500 m ² / g, preferably less than 350 m ² / g and even more preferably less than 250 m ² / g .

The hydrogenation step f) of the process according to the invention is preferably carried out in one or more fixed bed reactor (s).

In the hydrogenation zone (16), the feedstock is brought into contact with the hydrogenation catalyst in the presence of hydrogen and at operating temperatures and pressures allowing the hydrogenation of the olefinic unsaturated compounds present in the feedstock.

The operating conditions of the hydrogenation step f) of the process according to the invention are advantageously as follows: the temperature within said hydrogenation zone (16) is between 100 and 180 ^° C. and preferably between 120 and 165 ° C, the total pressure is between 0.5 and 6 MPa, preferably between 1 and 5 MPa and even more preferably between 2 and 5 MPa. The charge flow rate is such that the hourly volume velocity (ratio of the hourly volume flow rate at 15 ° C. of fresh liquid feedstock to the loaded catalyst volume) is between 1 and 50 h ^-1 , preferably between 2 and 20 h ^-1. and even more preferably between 4 and 20 h ^-1 . The hydrogen that feeds the hydrotreating zone is introduced at a rate such that the volume ratio hydrogen / hydrocarbons is between 5 to 300 Nl / l / h, preferably between 5 and 200, preferably between 10 and 150 Nl / l / h, and even more preferably between 10 and 50 Nl / l / / h.

Under these conditions, the olefinic type unsaturated compounds are hydrogenated more than 50%, preferably more than 75% and preferably more than 85%.

The effluent from step f) optionally undergoes a step of removing at least a portion of the water formed during the hydrogenation step f) and preferably all of the water formed, in a separator (37), shown in Figure 2.

This water may also contain a fraction of dissolved CO and CO2 from the Fischer Tropsch synthesis. This step takes place in the separator (37) and the water is removed by the pipe (39). It is also advantageous to add in the said step of removing at least a portion of the water nitrogen (N 2) or hydrogen (H 2) type gas in order to remove more CO and

CO2 dissolved by stripping.

At the end of step f) of the process according to the invention, at least a portion and preferably all of the liquid hydrogenated effluent is sent to a hydrocracking / hydroisomerization zone (19).

Step g)

According to step g) of the process according to the invention, at least a portion and preferably all of the liquid hydrogenated effluent from the hydrogenation step f) of the process according to the invention is sent, in the hydroisomerization / hydrocracking zone (19) containing the hydroisomerization / hydrocracking catalyst and preferably at the same time as a hydrogen flow.

The operating conditions in which the hydroisomerization / hydrocracking step g) of the process according to the invention is carried out are preferably as follows: The pressure is generally maintained between 0.2 and 15 MPa and preferably between 0.5 and 10 MPa and advantageously from 1 to 9 MPa, the space velocity is generally between 0.1 h -1 and 10 h -1 and preferably between 0.2 and 7 h -1 is preferably between 0.5 and 5.0 h -1. The hydrogen content is generally between 100 and 2000 normal liters of hydrogen per liter of filler and per hour and preferably between 150 and 1500 liters of hydrogen per liter of filler.

The temperature employed in this step is generally between 200 and 450 ⁰ C and preferably from 250 ⁰ C to 450 ⁰ C, advantageously 300 to 450 ⁰ C, and even more preferably greater than 320 ° C or, for example between 320-420 ⁰ vs.

Step g) of hydroisomerization and hydrocracking of the process according to the invention is advantageously carried out under conditions such that the pass conversion into products with boiling points greater than or equal to 37O ⁰ C into products having points. boiling point below 37O ⁰ C is greater than 80% by weight, and even more preferably at least 85%, preferably greater than 88%, so as to obtain middle distillates (gas oil and kerosene) having sufficiently good cold (pour point, freezing point) to meet the specifications in force for this type of fuel.

The catalysts of hydroxisomerization / hydrocracking

The majority of catalysts currently used in hydroisomerization / hydrocracking are of the bifunctional type associating an acid function with a hydrogenating function. The acid function is generally provided by supports with large surface areas (150 to 800 m2.g-1 generally) having a surface acidity, such as halogenated aluminas (chlorinated or fluorinated in particular), phosphorus aluminas, combinations of oxides of boron and aluminum, silica aluminas. The hydrogenating function is generally provided either by one or more metals of group VIII of the periodic table of the elements, such as iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum, or by an association of at least a group VI metal such as chromium, molybdenum and tungsten and at least one Group VIII metal.

In the case of bi-functional catalysts, the equilibrium between the two acid and hydrogenating functions is the fundamental parameter which governs the activity and the selectivity of the catalyst. A weak acidic function and a strong hydrogenating function give catalysts which are not very active and selective towards isomerization whereas a strong acid function and a low hydrogenating function give very active and cracking-selective catalysts. A third possibility is to use a strong acid function and a strong hydrogenating function in order to obtain a very active catalyst but also very selective towards isomerization. It is therefore possible, by judiciously choosing each of the functions to adjust the activity / selectivity couple of the catalyst.

Advantageously, the hydroisomerization-hydrocracking catalysts are bifunctional catalysts comprising an amorphous acid support (preferably a silica-alumina) and a hydro-dehydrogenating metal function preferably provided by at least one noble metal. The support is said to be amorphous, that is to say devoid of molecular sieves, and in particular of zeolite, as well as the catalyst. The amorphous acidic support is advantageously a silica-alumina but other supports are usable. When it is a silica-alumina, the catalyst preferably does not contain added halogen, other than that which could be introduced for the impregnation of the noble metal, for example.

More generally, and preferably, the catalyst does not contain added halogen, for example fluorine. In general, and preferably, the support has not been impregnated with a silicon compound.

According to a first preferred embodiment, the hydroisomerization / hydrocracking catalyst contains at least one hydrodehydrogenating element chosen from noble metals of group VIII, preferably platinum and / or palladium and at least one amorphous refractory oxide support, preferably silica alumina.

A preferred hydroisomerization / hydrocracking catalyst used in step g) of the process according to the invention comprises up to 3% by weight of metal of at least one hydrodehydrogenating element chosen from noble metals of group VIII, preferably deposited on the support, and very preferably, the noble metal of group VIII being platinum and a support comprising (or preferably consisting of) at least one silica-alumina, said silica-alumina having the following characteristics: - a weight content of silica SiO ₂ between 5 and 95%, preferably between 10 and 80%, more preferably between 20 and 60% and even more preferably between 30 and 50% by weight. an Na content of less than 300 ppm by weight and preferably less than 200 ppm by weight; a total pore volume of between 0.45 and 1.2 ml / g, measured by mercury porosimetry, the porosity of said silica-alumina being as follows:

The volume of the mesopores whose diameter is between 40 A and 150 A, and whose mean diameter varies between 80 and 140 A and preferably between 80 and 120 A, represents between 20 and 80% of the total pore volume measured by porosimetry at mercury, ii / The volume of macropores, whose diameter is greater than 500 A, and preferably between 1000 A and 10000 A represents between 20 and 80% of the total pore volume measured by mercury porosimetry,

a specific surface area of between 100 and 550 m ² / g, preferably between 150 and 500 m ² / g, preferably less than 350 m ² / g and even more preferably less than 250 m ² / g .

A second preferred hydroisomerization / hydrocracking catalyst used in step g) of the process according to the invention comprises up to 3% by weight of metal of at least one hydro-dehydrogenating element chosen from the noble metals of group VIII of the periodic classification and preferably the noble metal of group VIII being platinum, from 0.01 to 5.5% by weight of oxide of a doping element selected from phosphorus, boron and silicon and a non-zeolitic support to silica-alumina base containing an amount greater than 15% by weight and less than or equal to 95% by weight of silica (SiO ₂ ), said silica-alumina having the following characteristics: a mean porous diameter, measured by mercury porosimetry, between 20 and 140 A, a total pore volume, measured by mercury porosimetry, of between 0.1 ml / g and 0.5 ml / g, a total pore volume, measured by nitrogen porosimetry, of between 0.1 ml / g and 0.6 ml / g, a sp surface BET plastic between 100 and 550 m ² / g, a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 140 A less than 0.1 ml / g, a pore volume, measured by mercury porosimetry , contained in pores with a diameter of greater than 160 A less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, contained in pores with a diameter greater than 200 A, less than 0.1 ml / g, a porous volume, measured by mercury porosimetry, included in pores with diameters greater than 500 A less than 0.1 ml / g. an X-ray diffraction pattern which contains at least the principal characteristic lines of at least one of the transition aluminas included in the group consisting of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina. a packed packing density of the catalysts greater than 0.55 g / cm ³ . Advantageously, the characteristics associated with the corresponding catalyst are identical to those of the silica alumina described above.

The two stages f) and g) of the process according to the invention, hydrogenation and hydroisomerization-hydrocracking, can advantageously be carried out on the two types of catalysts in two or more different reactors, and / or in the same reactor.

According to a second preferred embodiment, the hydroisomerization / hydrocracking catalyst contains at least one hydrodehydrogenating element chosen from Group VIII non-noble metals and Group VIB metals and at least one amorphous refractory oxide support, preferably silica-alumina. .

Preferably, the group VIII metal is selected from nickel and cobalt, and the group VIB metal is selected from molybdenum and tungsten. Preferably, said catalyst is in sulphide form.

A third preferred hydroisomerization / hydrocracking catalyst used in step g) of the process according to the invention comprises at least one hydro-dehydrogenating element chosen from the non-noble metals of group VIII and the metals of group VIB of the periodic table, preferably between 2.5 and 5% by weight of Group VIII non-noble elemental oxide and between 20 and 35% by weight of Group VIB element oxide with respect to the weight of the final catalyst and, preferably, the non-noble metal of group VIII is nickel and the metal of group VIB is tungsten, optionally from 0.01 to 5.5% by weight of oxide of a doping element chosen from phosphorus, boron and silicon and preferably, from 0.01 to 2.5% by weight of oxide of a doping element and a non-zeolitic support based on silica-alumina containing an amount greater than 15% by weight and less than or equal to 95% by weight of silica (SiO _2), preferably an amount greater than 15% by weight and less than or equal to 50% by weight of silica, said silica-alumina having the following characteristics: a mean pore diameter, measured by mercury porosimetry, of between 20 and 140 A, a total pore volume, measured by mercury porosimetry, between 0.1 ml / g and 0.5 ml / g, a total pore volume, measured by nitrogen porosimetry, between 0.1 ml / g and 0.6 ml / g, a BET surface area of between 100 and 550 m ² / g, a pore volume, measured by mercury porosimetry, included in pores with a diameter greater than 140 A less than 0.1 ml / g, a volume Porous, measured by mercury porosimetry, included in pores with a diameter greater than 160 Λ less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, included in pores with diameters greater than 200 Å, less than 0.1 ml / g, a pore volume, measured by mercury porosimetry, included in pores with diameters greater than 500 A less than 0.1 ml / g. an X-ray diffraction pattern which contains at least the principal characteristic lines of at least one of the transition aluminas included in the group consisting of alpha, rho, chi, eta, gamma, kappa, theta and delta alumina. a packed packing density of the catalysts greater than 0.55 g / cm ³ .

Advantageously, the characteristics associated with the corresponding catalyst are identical to those of the silica alumina described above.

When the third preferred hydroisomerization / hydrocracking catalyst is used in step g) of the process according to the invention, said catalyst is sulphurized.

According to a first preferred embodiment of the process according to the invention, a catalyst containing palladium is used in stage e) of hydrogenation and in stage g) of hydroisomerization / hydrocracking a catalyst containing platinum.

According to a second preferred embodiment of the process according to the invention, a palladium-containing catalyst is used in the hydrogenation step f) and in the hydroisomerization / hydrocracking step g), a sulphurized catalyst containing at least one hydro-dehydrogenating element selected from Group VIII non-noble metals and Group VIB metals.

According to a third preferred embodiment of the process according to the invention, in the hydrogenation stage f), a catalyst containing at least one non-noble group VIII hydroxide dehydrogenating element and in the hydroisomerization stage g) is used. hydrocracking, a sulphurized catalyst containing at least one hydro-dehydrogenating element chosen from Group VIII non-noble metals and Group VIB metals. Step h)

According to step h) of the process according to the invention, the effluent from step g) (line 20) undergoes separation of unreacted hydrogen and light gases in a gas / liquid separator (33) and then recycling in step g) hydroisomerization / hydrocracking unreacted hydrogen and said light gases (line 22).

Said light gases include light C1-C4 gases, carbon monoxide (CO) and carbon dioxide (CO2) and water in vapor form.

The separation of said gases from the liquid effluent is carried out by one or more flashes, ie one or more balloons carrying out a separation of the gases and liquids introduced via the pipe (20), at staged temperatures and pressures to increase the recovery of hydrogen. This flash stage may advantageously be accompanied by a heat exchanger for recovering thermal energy and / or cooling the effluents of the separator flasks in order to minimize the losses of hydrogen. The process according to the invention makes it possible, by the use of ion exchange resin upstream of the hydrotreatment and hydrocracking steps, to reduce the total oxygen content of the feedstock and thus to limit the formation of carbon monoxide (CO) from the decomposition of the oxygenates present in the feed in the hydroisomerization / hydrocracking section. In fact, carbon monoxide (CO) is an inhibitor of the metal compounds present on the hydroisomerization / hydrocracking catalyst, and its content must be minimized so as not to require an increase in temperature to compensate for the drop in activity and maintain the conversion.

However, in the case where the gaseous effluent (22) resulting from said separation contains a high CO fraction, that is to say more than 10 ppm by volume, said gaseous effluent is advantageously sent to a methanation reactor (40), represented in FIG. 2, in which the conversion of CO and hydrogen into methane is advantageously carried out. The principle of methanation, the catalysts used are known to those skilled in the art and their use for purifying effluents containing H2 and CO is known. A purge is advantageously implemented (line 23) so as to eliminate the products formed during the methanation step (33). An additional hydrogen (pipe 24) is then advantageously made to compensate for this purge.

Step i)

The effluent from step h) for separating unreacted hydrogen and light gases from the process according to the invention is sent, according to step i) of the process according to the invention, in a train of distillation (26) via the pipe (21), which incorporates a distillation and optionally vacuum distillation, which aims to separate the conversion products of boiling point below 340 ⁰ C and preferably below 370 ⁰ C and including including those formed in step g) in the hydroisomerization / hydrocracking reactor (19), and to separate the residual fraction whose initial boiling point is generally greater than at least 340 ^° C. and preferably greater than or equal to at least 370 ^° C. Among the conversion products and hydroisomerized, it is separated in addition to the light gases C1-C4 (line 27) at least a gasoline fraction (or naphtha) (line 28), and at least a middle distillate fraction kerosene (line 29) and diesel (line 30). Preferably, the residual fraction, whose initial boiling point is generally greater than at least 340 ^° C. and preferably greater than or equal to at least 37 ^° C., is recycled (line 17) in step g) of the process according to the invention at the head of the zone (19) of hydroisomerization and hydrocracking. According to another embodiment of step i) of the process according to the invention, said residual fraction can provide excellent bases for the oils.

It may also be advantageous to recycle (line 32) at least partly and preferably entirely, in step g) (zone 19) at least one of the kerosene and diesel fuel cuts thus obtained. The gas oil and kerosene cuts are preferably recovered separately or mixed, but the cutting points are adjusted by the operator according to his needs. It has been found that it is advantageous to recycle a portion of the kerosene to improve its cold properties.

The products obtained

The gas oil (s) obtained has a pour point of at most 0 ^° C., generally below -10 ^° C. and often below -15 ^° C. The cetane number is greater than 60, generally greater than 65, often greater than 70.

The kerosene (s) obtained has a freezing point of not more than -35 ° C, generally less than -40 ^° C. The smoke point is greater than 25 mm, generally greater than 30 mm. In this process, the production of gasoline (not sought) is as low as possible. The yield of gasoline is always less than 50% by weight, preferably less than 40% by weight, advantageously less than 30% by weight, or 20% by weight or even 15% by weight.

EXAMPLE

Example 1: Implementation of the Process According to the Invention Step a)

The effluent from the Fischer-Tropsch synthesis unit is, at the output of the Fischer-Tropsch synthesis unit divided into two fractions, a light fraction, called cold condensate and a heavy fraction, called waxes.

The light fraction, called cold condensate, is fractionated into a C4- gas fraction boiling at a temperature below 15 ⁰ C and an intermediate fraction having an initial boiling point between 15 and 40 ⁰ C and a boiling point final temperature of between 300 and 450 ° C. The characteristics of the different fractions are given in Table 1 below:

Table 1: Composition of the different fractions

Step b)

The intermediate fraction C5-C30 passes over a commercial ion exchange resin Amberlyst 35 sold by the company Rohm and Haas, said resin simultaneously allowing the capture of dissolved metals in the feedstock and the esterification of alcohols and carboxylic acids to ester . Said resin consists of copolymers of di-vinyl benzene and polystyrene having a degree of crosslinking of 20 and an acidic strength, assayed by potentiometry during neutralization with a KOH solution, of 4.15 mmol H + equivalent / g.

Step b) is carried out at a temperature of 100 ^° C., a pressure of 1 MPa, and a hourly volume velocity of 1 ^{h -1} . Under these conditions, 95% of the acids are esterified, the analysis of the acid conversion being given by difference in potash titration between the feedstock and the effluent according to the ASTM D 664 method. The composition of the outlet effluent is given in Table 2.

Table 2: composition of the effluent (7) from step b) after esterification

Step cl

The effluent from step c) then undergoes elimination of the water formed during said step b) by decantation / coalescence in suitable equipment known to those skilled in the art. The effluent at the outlet contains 300 ppm of water.

Step d)

The heavy fraction, called waxes, from the Fischer-Tropsch synthesis unit (line 11) is passed over a guard bed composed of ACT 961 sold by Axens at a temperature of 80 ^° C. and a pressure of 1 MPa. , a liquid space velocity of 1 h-1. At the outlet, the effluent (line 13) contains less than 1 ppm of fines.

Step e)

The purified intermediate fraction (line 10) and the wax fraction from step d) (line 13) are mixed in totality. The composition of the mixture thus produced is given in Table 3 below:

Table 3: Composition of the recombined fraction

Fi step

The entire effluent from step e) then undergoes a hydrogenation step in the presence of hydrogen and the hydrogenation catalyst of commercial name LD265 marketed by the company Axens, said catalyst comprising 0.3% by weight of palladium deposited on a surface alumina of 69 m2 / g.

The hydrogenation step f) is carried out at a reaction temperature of 130 ^° C. at a pressure of 3.5 MPa, the hydrogen is introduced at a rate such that the hydrogen / hydrocarbon volume ratio is 32 Nl. / l / h and at a speed hourly volume of 10 h-1. Under these conditions, the conversion to olefins is 85% by weight.

The liquid effluent from the hydrogenation step f) has the composition described in Table 4:

Table 4: Composition of the effluent from step f)

Step g)

All of the effluent from the hydrogenation step f) undergoes a hydroisomerization / hydrocracking step, in the presence of fresh hydrogen and a hydroisoméhation / hydrocracking catalyst, in which the residual fraction of initial boiling point above 370 ⁰ C and unreacted hydrogen and light gases. The hydroisomerization / hydrocracking catalyst comprises 0.6% by weight of platinum and a support comprising 29.3% by weight of silica SiO ₂ and 70.7% by weight of Al ₂ O ₃ alumina, an Na content of 100 ppm by weight. a total pore volume of 0.69 ml / g measured by mercury porosimetry, a volume of mesopores with an average diameter of 80 A and representing 78% of the total pore volume measured by mercury porosimetry, a volume of macropores, whose diameter is greater than 500 A ₁ representing 22% of the total pore volume measured by mercury porosimetry, and a specific surface area of 300 m ² / g.

The hydroisomerization / hydrocracking step is carried out under the conditions described in Table 5.

The pass conversion to products with a boiling point greater than or equal to 370 ° C. in products with a boiling point below 37 ^° C. is 85%.

Table 5: Operating Conditions of the Hydroisomerization / Hydrocracking Step

Step h)

The effluent from the hydroisomerization / hydrocracking stage undergoes separation in a gas / liquid separator, unreacted hydrogen and light gases, which are recycled in the hydroisomerization / hydrocracking step, so that recover a liquid effluent. The Carbon monoxide (CO) content generated by passes in the gaseous effluent is limited to 1; 1% by weight.

Step i)

The liquid effluent resulting from the separation step h) is then sent to a distillation train so as to separate the light products formed during these steps: the gases (C1-C4), a gasoline cut, a diesel fuel cut and a kerosene cut, and also a fraction, called residual fraction, which has an initial boiling point equal to 370 ⁰ C which is recycled in full at the inlet of the hydroisomerisation / hydrocracking reactor in order to maximize the production of diesel fuel and kerosene. The yields are given in Table 6.

Example 2: Comparative

A process for producing middle distillates from the same paraffinic effluent from the Fischer tropsch unit and divided into two fractions, a heavy fraction called waxes and a light fraction called cold condensate, as used in Example 1 and comprising a hydrogenation step followed by a hydroisomerization / hydrocracking step, without prior step of passing on at least one ion exchange resin is carried out for comparison.

The light fraction, called cold condensate, is fractionated into a C4-gas fraction boiling at a temperature below 15 ^° C. and an intermediate fraction having an initial boiling point of between 15 and 40 ^° C. and a boiling point. final between 300 and 450 ⁰ C. the characteristics of the different fractions are given in table 1 of example 1.

Decontamination step of the heavy fraction called wax

The heavy fraction, called waxes, from the Fischer-Tropsch synthesis unit (line 11) is passed over a guard bed composed of ACT 961 sold by Axens at a temperature of 80 ^° C. and a pressure of 1 MPa. , a liquid space velocity of 1 h-1. At the outlet, the effluent (line 13) contains less than 1 ppm of fines. Recombination step of the fractions

Said intermediate fraction, not previously purified by passing on at least one ion exchange resin and the wax fraction from the above decontamination step are mixed in all. The composition of the mixture thus produced is given in Table 7:

Hydrogenation step The recombinant C5 + paraffinic fraction undergoes a hydrogenation step in the presence of hydrogen and the LD265 commercial name hydrogenation catalyst sold by the company Axens, said catalyst comprising 0.3% by weight of palladium deposited on an alumina specific surface area 69 m2 / g.

To keep a conversion to olefin of 85 weight%, as in Example 1, the hydrogenation step is carried out at a reaction temperature of 150 ⁰ C at a pressure of 3.5 MPa, the hydrogen is introduced at a rate such that the volume ratio hydrogen / hydrocarbon is 32 Nl / l / h and at an hourly volume velocity of 8 h-1. Under these conditions, the conversion to olefins is maintained at 85% by weight.

The liquid effluent resulting from the hydrogenation stage has the composition described in Table 8:

Table 8: Com osition of the effluent from the hydration stage

Hydroisomerization / Hydrocracking Step All of the effluent from the hydrogenation step undergoes a hydroisomerization / hydrocracking step, in the presence of fresh hydrogen and a hydroisomerization / hydrocracking catalyst, in which are recycled the residual fraction of initial boiling point greater than 37O ⁰ C and unreacted hydrogen and light gases. The hydroisomerization / hydrocracking catalyst comprises 0.3% by weight of platinum and a support comprising 29.3% by weight of silica SiO ₂ and 70.7% by weight of alumina Al ₂ O ₃ , a content in Na of 100 ppm by weight, a total pore volume of 0.69 ml / g measured by mercury porosimetry, a volume of mesopores with a mean diameter of 80 A and representing 78% of the total pore volume measured by mercury porosimetry , a volume of macropores, whose diameter is greater than 500 A, representing 22% of the total pore volume measured by mercury porosimetry, and a specific surface area of 300 m ² / g.

The hydroisomerization / hydrocracking step is carried out under the conditions described in Table 9.

To maintain the conversion per pass of products boiling point higher than or equal to 370 ⁰ C in products with boiling points less than 370 C ^c 85%, the temperature is raised and adjusted to 37o C. ⁰

Table 9: Operating conditions of the isomerization / hydrocreation stage

Separation step

The effluent resulting from the hydroisomerization / hydrocracking step is then separated in a gas / liquid separator, unreacted hydrogen and light gases, which are recycled in the hydroisomerization / hydrocracking step, so that recovering a liquid effluent. The content of carbon monoxide (CO) generated by passes into the gaseous effluent is 2% by weight.

The presence of carbon monoxide (CO) resulting from the decomposition of the oxygenated compounds in the hydrocracking section, not removed by passage on at least one ion exchange resin prior to hydrogenation and hydroisomerization / hydrocracking and being an inhibitor The activity of the hydroisomerization / hydrocracking catalyst requires to maintain the conversion an increase in the reaction temperature.

Final distillation stage

The liquid effluent from the separation step is then sent to a distillation train so as to separate the light products formed during these steps: the gases (C1-C4), a gasoline cut, a diesel cut and a cut kerosene, and also a fraction, called residual fraction, which has an initial boiling point equal to 370 ⁰ C which is recycled in all at the inlet of the hydroisomerization / hydrocracking reactor to maximize the production of diesel and kerosene.

The yields are given in Table 10.

Table 10: Yields of Different Seeds at Multiple Separation

The presence of carbon monoxide (CO) in the gaseous effluent inhibiting the hydrogenating function of the hydrocracking catalyst modifies not only the activity but also the selectivity of said catalyst in middle distillates. It is found that the absence of a prior step of passing on at least one ion exchange resin requiring an increase in temperature to maintain the conversion causes an increase in the production of undesirable light gas and naphtha fraction by over cracking.

The process according to the invention exemplified in Example 1 thus makes it possible, by the use of ion exchange resin upstream of the hydrotreatment and hydroisomerization / hydrocracking steps, to reduce the total oxygen content of the feedstock. and thus to limit the formation of carbon monoxide (CO) from the decomposition of the oxygenates present in the feed in the hydroisomerization / hydrocracking section. In fact, carbon monoxide (CO) is an inhibitor of the metal compounds present on the hydroisomerization / hydrocracking catalyst, and its content must be minimized so as not to require an increase in temperature to compensate for the decrease in activity and maintain the conversion. It is therefore found that the process according to the invention allows a reduction in the production of carbon monoxide (CO) (1.1% by weight) by the implementation of step b) according to the invention compared to a process not according to the invention does not implement said step of passing on at least one ion exchange resin and allows a decrease in the temperatures used in the hydrogenation step and hydroisomerization / hydrocracking to obtain the same conversion of 85% to products with a boiling point greater than or equal to 370 ^° C. in products with boiling points below 370 ° C.

Claims

1. Process for producing middle distillates from a paraffinic feedstock produced by Fischer-Tropsch synthesis and divided into two fractions, a light fraction, called

S cold condensate, and a heavy fraction, called waxes, comprising the following steps: a) fractionation of said light fraction, called cold condensate into two fractions, a gaseous fraction C4- boiling at a temperature below 15 ° C and a intermediate fraction having an initial boiling point between 15 and 40 ^° C. and a final boiling point of between 300 and 450 ^° C., b) passing on at least one ion exchange resin of said intermediate fraction, at a temperature between 50 ⁰ C and 150 "C ₁ at a total pressure between 0.7 and 2.5 MPa, at an hourly space velocity between 0.2 and 2.5 h-1, c) removal of at least a portion of the water of the effluent from step b), d) decontamination by passing over a first guard bed containing at least one guard bed catalyst of said heavy fraction called waxes, e) recombination of the purified intermediate fraction resulting from step c ) and the effluent from step d) to obtain a purified C5 + fraction, f) hydrogenation of the olefinic unsaturated compounds of at least a portion of the purified C5 + fraction from step e) in the presence of hydrogen and a hydrogenation catalyst, g) hydroisomerization / hydrocracking in the presence of hydrogen and a hydroisomerization / hydrocracking catalyst of at least a portion of the effluent from step f) h) separation and recycling in step g) hydroisomerization / hydrocracking unreacted hydrogen and light gases, i) distillation of the effluent from step h).

2. The method of claim 1 wherein prior to step b) said intermediate fraction undergoes a step of decontamination by passing on a guard bed containing at least a guard bed catalyst.

3. Method according to one of claims 1 or 2 wherein said C5 + liquid paraffinic fraction from step b) passes over a single ion exchange resin for simultaneously performing the esterification of alcohols and carboxylic acids into ester and the capture of dissolved metals in the charge.

4. Process according to claim 3, in which said resin is used at a temperature of between 80 ^° C. and 150 ^° C., at a pressure of between 1 and 2 MPa and at an hourly space velocity of between 0.5 and 1. , 5 h-1.

5. Method according to one of claims 3 or 4 wherein said resin is constituted by copolymers of di-vinyl benzene and polystyrene having a degree of crosslinking of between 20 and 35% and an acidic strength, determined by potentiometry during neutralization with a KOH solution of 0.2 to 6 mmol H + equivalent / g.

6. Method according to one of claims 3 or 4 wherein said resin is a polysiloxane grafted with alkylsulfonic acid groups (type -CH2-CH2-CH2-SO3H), size between 0.5 and 1, 2 mm and potentiometric strength assayed by potentiometry during neutralization with a KOH solution, 0.4 to 1.5 mmol H + equivalent / g.

7. Method according to one of claims 1 or 2 wherein said C5 + liquid paraffinic fraction from step b) passes on two different ion exchange resins different in two different reactors.

8. The method of claim 7 wherein the reactor containing the ion exchange resin for the capture of metals is carried out upstream of the reactor containing the ion exchange resin for the esterification of alcohols and carboxylic acids.

9. Method according to one of claims 7 or 8 wherein said first resin is a resin consisting of copolymers of di-vinyl benzene and polystyrene having a degree of crosslinking of between 1 and 20% and an acid strength dosed by potentiometry during neutralization with a KOH solution, between 1 and 15 mmol H + equivalent / g.

10. Method according to one of claims 7 to 9 wherein said first resin is carried out at a temperature between 50 ⁰ C and 110 ⁰ C at a pressure between

1 and 2 MPa and at an hourly space velocity of between 0.2 and 1.5 h -1.

11. Method according to one of claims 1 to 10 wherein said guard bed catalyst used without step d) and in the optional decontamination step prior to step b) comprises a macroporous mercury volume for a mean diameter at 50 nm greater than 0.1 cm3 / g, and a total volume greater than 0.60 cm3 / g.

12. Method according to one of claims 1 to 11 wherein the effluent from step d) passes, before the recombination step e), on at least one ion exchange resin at a temperature between 80 ° ⁰ C and 150 ⁰ C, at a total pressure of between 0.7 and 2.5 MPa, at an hourly space velocity of between 0.2 and 2.5 h -1.

13. Method according to one of claims 1 to 11 wherein the paraffinic filler produced by Fischer-Tropsch synthesis is produced from a synthesis gas produced from natural gas according to the GTL route also called in the English terminology. Saxon gas-to-liquid.

14. Method according to one of claims 1 to 11 wherein the paraffinic filler produced by Fischer-Tropsch synthesis is produced from a synthesis gas produced from coal according to the CTL route also called in the English terminology. coal-to-liquid.

15. Method according to one of claims 1 to 11 wherein the paraffinic feedstock produced by Fischer-Tropsch synthesis is produced from a synthesis gas produced from biomass according to the BTL pathway also called in the English terminology. biomass-to-liquid.