EP1612255B2 - Hydrodesulfurization process for naphtha fractions using a catalyst with controlled porosity. - Google Patents

Description

The present invention relates to a desulfurization process using a catalyst containing at least one support, and an active phase comprising for example a metal. The process allows the hydrodesulphurization of gasolines, and especially the gasolines resulting from a catalytic cracking process (FCC Fluid Catalytic Cracking according to English terminology, that is to say catalytic cracking in a fluidized bed).

In particular, the production of reformulated species that meet the new environmental standards requires a significant reduction in their sulfur content. Thus, the current and future environmental standards require refiners to lower the sulfur content in the gasoline pool to values of 50 ppm or less in 2005 and 10 ppm in the European community by 1 January 2009. The feedstock to be treated is usually a sulfur-containing gasoline cut such as for example a cut from a coking unit, visbreaking, steam cracking or catalytic cracking (FCC). Said feed is preferably constituted by a gasoline cut from a catalytic cracking unit whose boiling point range typically ranges from that of hydrocarbons with 5 carbon atoms to about 250 ° C. This gasoline can possibly be composed of a significant fraction of gasoline coming from other production processes such as atmospheric distillation (generally called straight run fuel by the refiner) or conversion processes (coker or steam-cracking gasoline).

Catalytic cracking gasolines, which can constitute 30 to 50% by volume of the gasoline pool, have high olefin and sulfur contents. The sulfur present in the reformulated gasoline is attributable, to nearly 90%, to the gasoline resulting from catalytic cracking. Desulphurisation of species and mainly FCC species is therefore of obvious importance for achieving standards. Hydrotreating or hydrodesulfurization of catalytic cracking gasoline, when it is carried out under the conditions conventionally known to those skilled in the art, makes it possible to reduce the sulfur content of the cut. However, this process has the major disadvantage of causing a very large drop in the octane number of the cut, because of the hydrogenation or saturation of a large part or even all of the olefins under the conditions of the hydrotreatment. It has therefore been proposed methods for deep desulfurization of FCC gasoline while maintaining the octane number to an acceptable level. Thus, the patent US 5,318,690 proposes a process consisting in splitting the gasoline, softening the light fraction and hydrotraying the heavy fraction on a conventional catalyst and then treating it on a zeolite ZSM5 to recover the initial octane. The patent WO 01/40409 claims the treatment of an FCC gasoline under conditions of high temperature, low pressure and high hydrogen / charge ratio. Under these particular conditions, the recombination reactions involving H ₂ S formed by the desulphurization reaction and the olefins to give rise to mercaptan formation are minimized.

The improvement of the reaction selectivity (hydrodesulfurization / hydrogenation) sought can therefore be obtained by the choice of the process but in all cases the use of an inherently selective catalytic system is essential. In general, the catalysts used for this type of application are sulphide catalysts containing a group VIB element (Cr, Mo, W) and a group VIII element (Fe, Ru, Os, Co, Rh , Ir, Pd, Ni, Pt).

Obtaining selective catalysts for the selective hydrodesulphurization of olefinic gasoline cuts is taught in many patents. Some patents propose the use of supports other than the alumina support conventionally used for hydrotreating catalysts, such as, for example, magnesia-based supports (patent US 4,203,829 ; patent US 4,140,626 ), spinel (patent US5,525,211 ), carbon (patent US 5,770,046 ), hydrotalcite (patent 5340466 ). Other patents claim the use of a controlled mesoporosity catalyst such as the patent US 6,013,598 which claims the use of a median porous diameter catalyst (measured by mercury porosimetry) of between 7.5 and 17.5 nm. Despite the progress made, the search for new improved selectivity catalysts remains an important goal in the field of hydrotreating of cracked species.

• une hydrogénation des oléfines limitée à fort taux de désulfuration,
• une bonne stabilité du système catalytique et une opération continue pendant plusieurs années.

To be competitive, hydrodesulfurization processes must meet two main constraints that are:

• a hydrogenation of olefins limited to high desulphurization rate,
• good stability of the catalytic system and continuous operation for several years.

• les composés soufrés insaturés que sont le thiophène, les méthyl-thiophènes, les diméthyl-thiophènes, éthyl-thiophènes, et autres alkylthiophènes, les benzothiophènes et alkylbenzothiophènes.
• les composés soufrés saturés que sont les mercaptans, les sulfures cycliques ou aliphatiques, les disulfures.

In addition, to perform a deep desulfurization, it is necessary to treat all the sulfur compounds present in cracking gasoline and in this context the catalytic cracking gasolines can be classified into two families:

• unsaturated sulfur compounds such as thiophene, methyl-thiophenes, dimethyl-thiophenes, ethyl-thiophenes, and other alkylthiophenes, benzothiophenes and alkylbenzothiophenes.
• saturated sulfur compounds such as mercaptans, cyclic or aliphatic sulfides, disulfides.

The residual sulfur compounds present in the desulfurized gasolines by deep hydrodesulfurization comprise so-called recombinant mercaptans, resulting from the addition of the H ₂ S formed during the reaction, on the olefins present, as well as unsaturated sulfur compounds such as thiophene and alkylthiophenes. The presence of the so-called recombination mercaptans explains at least in part that when it is desired to desulphurize in depth gasolines comprising an olefin fraction, a very high growth rate of hydrogenation of olefins is observed for the high levels of desulfurization. . Thus, when the desired desulfurization rate approaches 100%, the saturation rate of olefins increases very significantly. The use of more selective catalysts can allow when desulfurization rates close to 100% are desired, however, to limit the hydrogenation of olefins, see the formation of recombinant mercaptans. One of the main issues of deep desulphurization is thus to find methods for achieving high selectivities that is to say to minimize the hydrogenation reactions of olefins while treating the residual sulfur compounds such as mercaptans.

Among the possible solutions for achieving the desulfurization rates imposed by current or future standards, it may be advantageous to use a desulfurization process in at least two stages.

Thus, the patent application EP 1 031 622 A1 discloses a process for desulfurizing olefinic gasolines comprising at least two steps, a step of hydrogenation of unsaturated sulfur compounds and a step of decomposition of saturated sulfur compounds. As described in the patent, the invention is based on a two-step sequence such that the first step allows the elimination of unsaturated sulfur compounds from saturated sulfur compounds and the second step decomposes the saturated sulfur compounds into H ₂ S with hydrogenation of olefins limited.

The patent US6,231,753 describes a process for the hydrodesulfurization of olefinic gasolines comprising a first hydrodesulphurization step, an H2S extraction step and a second hydrodesulphurization step, the overall desulfurization rate and the temperature of this second step being greater than those from the first.

The patent US6,231,754 describes a process in which a spent hydrotreatment catalyst is subsequently used in a higher temperature hydrodesulfurization step. The pore diameters of the catalyst are described as between 6 and 20 nm and the surface concentration of MoO ₃ between 0.5. 10 ^-4 and 3.10 ^-4 g / m ² .

Requirement WO 03/099963 discloses a two-step process wherein the second step is carried out with a catalyst less loaded with metals and having a pore diameter equal to or greater than the catalyst used in the first step. The average pore diameter of the catalysts is between 6 and 20 nm and the surface concentration of MoO ₃ is between 0.5. 10 ^-4 and 3.10 ^-4 g / m ² .

SUMMARY OF THE INVENTION

According to the present invention, there has been found a process for reducing the total sulfur content of hydrocarbon cuts and preferably FCC gasoline cuts, without loss of gasoline yield and minimizing the decrease of the octane number.

The hydrodesulfurization process of a gasoline according to the invention uses a catalyst comprising a support and an active phase comprising at least one metal characterized in that the average pore diameter of said catalyst is greater than 20 nanometers, preferably between 20 and 100 nm.

Preferably, the catalyst according to the invention contains at least one Group VI metal, more preferably it additionally contains at least one Group VIII metal. The surface density of the Group VI metal is preferably between 2 × 10 ^-4 and 40 × 10 ^-4 grams of oxide of said metal per m ² of support.

In the process according to the invention, the support is preferably chosen from the group consisting of aluminas, silica, silica alumina or titanium or magnesium oxides used alone or in admixture with alumina or silica-alumina. . More preferably, the support is at least in part constituted by an alumina. According to a variant of the invention, the specific surface area of the support is less than 200 m ² / g.

The hydrodesulfurization process according to the invention comprises at least two successive stages of hydrodesulfurization and a catalyst whose average pore diameter is greater than 20 nanometers is used in at least one of said stages. The successive steps are performed without intermediate degassing.

The process according to the invention comprises a succession of hydrodesulfurization steps and the activity of the catalyst of a step n + 1 is between 1% and 90% of the activity of the catalyst of step n, the catalyst of step n + 1 comprising a lower metal content than the catalyst of step n.

According to another embodiment of the process according to the invention, the reaction temperature of step n + 1 is greater than that of step n. In another embodiment, the catalyst of step n + 1 is the catalyst of step n which has undergone partial deactivation. In this case, for example, the deactivation of the catalyst can be achieved by contacting the catalyst with a feed containing a hydrocarbon fraction comprising olefins at a temperature at least 250 ° C. It is also possible to recycle the catalyst from step n to step n + 1 when its activity has decreased by at least 10%.

DETAILED DESCRIPTION OF THE INVENTION

The process according to the invention uses at least one hydrodesulfurization catalyst comprising at least one Group VI metal (M ^VI ) and / or at least one Group VIII metal (M ^VIII ) on a support. The Group VI metal is generally molybdenum or tungsten the Group VIII metal generally nickel or cobalt. The catalyst support is usually a porous solid selected from the group consisting of aluminas, silicon carbide, silica, silica-aluminas or titanium or magnesium oxides used alone or in admixture with the alumina or silica-alumina. It is preferably selected from the group consisting of silica, the family of transition aluminas and silica-aluminas. Very preferably, the support consists essentially of at least one transition alumina, that is to say it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight. % weight, or even at least 90% weight of transition alumina. It may optionally consist solely of a transition alumina.

The specific surface of the support is generally less than 200 m ² / g, most often less than 150 m ² / g. The porosity of the catalyst before sulfurization is such that it has an average pore diameter greater than 20 nm, preferably greater than 25 nm or even 30 nm and often between 20 and 140 nm, preferably between 20 and 100 nm, and very preferably between 25 and 80 nm. The pore diameter was measured by mercury porosimetry according to ASTM D4284-92 with a wetting angle of 140 °.

The surface density of the group VI metal is in the range according to the invention between 2.10 and 40.10 ^{-4 -4} gram of oxide of said metal per m ² of support, preferably between 4.10 ^-4 and 16.10 ^-4 g / m ^2.

According to the invention, the molar ratio M ^VIII / (M ^VI + M ^VIII ) is typically greater than 0.1, preferably between 0.2 and 0.6 and very preferably between 0.2 and 0. 5.

The catalyst according to the invention can be prepared using any technique known to those skilled in the art, and in particular by impregnation of the elements of groups VIII and VIB on the selected support. This impregnation may, for example, be carried out according to the method known to those skilled in the art in the dry-impregnation terminology, in which the quantity of desired elements in the form of soluble salts in the chosen solvent, for example demineralized water, so as to fill the porosity of the support as exactly as possible. The support thus filled with the solution is preferably dried. The preferred support is alumina which can be prepared from any type of precursors and shaping tools known to those skilled in the art.

After introduction of the elements of groups VIII and VIB, and possibly forming the catalyst, it undergoes an activation treatment. This treatment generally aims to transform the molecular precursors of the elements in the oxide phase. In this case it is an oxidizing treatment, but a direct reduction or even a simple drying of the catalyst can also be carried out. In the case of an oxidizing treatment, also known as calcination, this is generally carried out under air or under dilute oxygen, and the treatment temperature is generally between 200 ° C. and 550 ° C., preferably between 300 ° C. C and 500 ° C. In the case of a reducing treatment, this is generally carried out under pure hydrogen or preferably diluted, and the treatment temperature is generally between 200 ° C. and 600 ° C., preferably between 300 ° C. and 500 ° C. ° C.

Group VIB and VIII metal salts useful in the catalyst preparation process are, for example, cobalt nitrate, nickel nitrate, ammonium heptamolybdate or ammonium metatungstate. Any other salt known to those skilled in the art having sufficient and decomposable solubility during the activation treatment can also be used.

The catalyst is usually used in a sulfurized form obtained after treatment in temperature in contact with a decomposable organic sulfur compound and generating H ₂ S or directly in contact with a gas stream of H ₂ S diluted in H ₂ . This step can be carried out in situ or ex situ (inside or outside the reactor) of the hydrodesulfurization reactor at temperatures between 200 and 600 ° C and more preferably between 300 and 500 ° C.

The present invention also relates to a process for the desulphurization of gasolines comprising olefins comprising at least two hydrodesulphurization stages and intended to minimize both the content of compounds which are the most refractory to hydrodesulfurization, such as the thiophene compounds and the so-called recombination mercaptans, derived from the addition of H ₂ S to the olefins while limiting the degree of hydrogenation of the olefins associated with the removal of the sulfur compounds. At least one of the steps of the hydrodesulfurization process uses a catalyst as previously described.

The method comprises at least two steps. A hydrodesulfurization first stage A is preferably carried out in a fixed bed reactor, generally in the vapor phase, on any catalyst conventionally used for this application. The use of so-called "selective" catalysts is preferred because it makes it possible to limit the hydrogenation of the olefins while maximizing the hydrodesulfurization. This first step is followed by a second step B, for example without operation between steps A and B other than a warming of the effluent of step A. Step B is characterized by the fact that it is carried out on a catalyst having a catalytic activity in conversion of thiophene of between 1% and 90%, or even between 1% and 70% and preferably between 1% and 50% of the activity of the catalyst of stage A. The catalyst implemented in step B may be either a catalyst whose catalytic formulation has been optimized to achieve the desired catalytic activity, or a partially deactivated catalyst.

According to the invention, the use of catalysts which are preferably more selective in series makes it possible to limit the hydrogenation of the olefins to the high levels of desulfurization. he It has been observed that such a sequence makes it possible, by an inexpensive device, to significantly improve the selectivity of the desulfurization reaction by minimizing the olefin saturation rate while maintaining a degree of conversion of the sulfur compounds to H ₂ S Student. This device also has the advantage that it allows, for a diagram without extraction of the H ₂ S between the two reactors, to improve the selectivity of the process compared with a desulphurization performed in a single step. Compared to the teaching described in the application EP 1 031 622 A1 the implementation of the present process makes it possible to achieve higher desulphurization rates for the same degree of hydrogenation of the olefins because the unsaturated compounds not converted in the first step can be converted in the second step.

In the particular case where the catalyst of step B is the same catalyst as that of step A, but whose catalytic activity has been decreased by deactivation, the device is most often based on a set of at least two or three reactors and can be carried out as follows: the reactor of step A contains the fresh catalyst and the reactor of step B contains the spent catalyst. When the catalyst of step A is deactivated, the reactor containing the catalyst of deactivated step A is used in the second step, a reactor containing fresh catalyst is started and placed in step A. The reactor containing catalyst B is stopped the catalyst is replaced by fresh catalyst and the reactor is put on hold. This scheme makes it possible to operate the desulfurization unit without interruption during the replacement of the spent catalyst, while maximizing the selectivity of the process.

This embodiment is more particularly of interest for an operation of the low pressure and high temperature hydrodesulphurization section for the two stages, conditions for which the formation of the recombination mercaptans is minimized but which results in a rapid deactivation of the catalysts. hydrodesulfurization. By low pressure is meant relative pressures generally less than 2 relative MPa, and preferably less than 1.5 MPa relative or even 1 MPa relative and temperatures generally greater than 250 ° C or 260 ° C and most often greater than 280 ° C.

• un taux de désulfuration généralement inférieur à 98%, de façon préférée inférieur à 95% et de façon très préférée inférieure à 90%.
• un taux d'hydrogénation des oléfines inférieur à 60% et de façon préférée inférieur à 50%.

Stage A is generally characterized by:

• a desulfurization rate generally less than 98%, preferably less than 95% and very preferably less than 90%.
• a degree of hydrogenation of the olefins of less than 60% and preferably less than 50%.

• un taux de désulfuration généralement inférieur à 98%, de façon préférée inférieur à 95% et de façon très préférée inférieure à 90%.
• un taux d'hydrogénation des oléfines- inférieur à 60% et de façon préféré inférieur à 50%.
• une température opératoire supérieure à celle de l'étape A, de façon préférée supérieure de plus de 10°C à la température de l'étape A et de manière plus préférée supérieure de plus de 20°C à la température de l'étape A,
• l'utilisation d'un catalyseur dont l'activité par unité de volume mesurée en conversion du thiophène est comprise entre 1% et 90% de l'activité du catalyseur de l'étape A. Cette activité catalytique est mesurée par un test molécule modèle décrit dans la suite du texte.

Stage B is most often characterized by:

• a desulfurization rate generally less than 98%, preferably less than 95% and very preferably less than 90%.
• a degree of hydrogenation of the olefins of less than 60% and preferably less than 50%.
• a higher operating temperature than that of step A, preferably more than 10 ° C higher than the temperature of step A and more preferably more than 20 ° C above the temperature of step A ,
• the use of a catalyst whose activity per unit volume measured in conversion of thiophene is between 1% and 90% of the activity of the catalyst of step A. This catalytic activity is measured by a test molecule model described in the rest of the text.

The pressure of steps A and B is generally between about 0.4 MPa and 3 MPa, preferably between 0.6 MPa and 2.5 MPa, the flow rate of hydrogen is such that the ratio of hydrogen flow rates in normal liters per hour on the hydrocarbon flow rate in liters per hour is between 50 and 800 and preferably between 60 and 600. The temperature of step A is between 150 ° C and 450 ° C, preferably between 200 and 200 ° C. ° C and 400 ° C and more preferably between 230 ° C and 350 ° C and the temperature of step B is between 150 ° C and 450 ° C, preferably between 210 ° C and 410 ° C and more preferably between 240 ° C and 360 ° C.

Steps A and B are performed according to a preferred mode in sequence without intermediate additional step. It is therefore possible to implement them in the same reactor. In this case, the catalytic zone corresponding to step B will be operated at an average temperature above the minimum of 10 ° C to the catalytic zone corresponding to step A. This temperature difference can come either from the heat of reaction released by the hydrogenation of olefins, either by injection between the catalytic zones A and B of a hotter fluid selected from hydrogen or an inert gas such as nitrogen, the feed or the fluid from the recycling of a fraction of the effluent of the process.

Steps A and B can also be carried out in a catalytic column from which the gaseous compounds are withdrawn at the top under normal conditions of temperature and pressure. In this case, the catalytic zone of step A will be arranged higher in the column than the catalytic zone of step B.

The catalyst of step B differs from the catalyst of step A by a catalytic activity of between 1% and 90%, or even between 1 and 70% and preferably between 1% and 50% of the catalytic activity of the catalyst. Step A. The catalysts of Steps A and B are carried out in sulphurous form. The sulphurization procedure can be carried out in situ or ex situ by any sulphurization method known to those skilled in the art.

avec :

• A : activité du catalyseur, en min^-1. cm³ _catalyseur ^-1 (centimètre cube^-1)
• k constante de vitesse pour la conversion du thiophène, en min^-1
• m_catalyseur : masse de catalyseur utilisée en g
• DRT : densité de remplissage tassée du catalyseur, en cm³/g

The activity of the catalyst can be defined with respect to the rate constant for the conversion of normalized thiophene per volume of catalyst evaluated in a test on model molecules. The speed constant is calculated considering an order 1 for the reaction: $$\mathrm{AT}=\mathrm{k}/\left({\mathrm{m}}_{\mathrm{catalyst}}\times {\mathrm{DRT}}_{\mathrm{catalyst}}\right)$$

with:

• A: catalyst activity, in min ^-1 . cm ³ _catalyst ^-1 (cubic centimeter ^-1 )
• k rate constant for thiophene conversion, in min ^-1
• m _catalyst : mass of catalyst used in g
• DRT: packed packing density of catalyst, in cm ³ / g

When the catalyst used is a new catalyst prepared to exhibit reduced activity, it may be envisaged to prepare a new catalyst by impregnating a small amount of metals on the support. Typically, the contents of Group VIII and Group VIB metals deposited on the support will not exceed 10 and 14% by weight in oxide form and preferably 7.8% and 10% weight respectively in oxide form (to remain consistent with the maximum Co / Co + Mo ratio of 0.6 of the preferred range). The support used generally contains silica, silicon carbide, titanium oxide or magnesium oxide and / or alumina but will preferably be predominantly composed of alumina.

The catalyst of step B can also be a deactivated hydrotreatment catalyst. For example, a used catalyst from a distillate hydrodesulfurization unit or any other hydrodesulphurization process present in the refinery could be used, provided that the residual activity measured by the method described in Example 6 did not occur. does not exceed 90%, even 70% and preferably 50% of the activity of the catalyst of step A.

The catalyst of step B may finally have been deactivated by treatment of a cut comprising olefins. The spent catalysts generally have a decreased activity due to the presence of carbon deposition due to the polymerization of the hydrocarbons treated on the catalyst.

The present invention can be implemented as follows: The gasoline to be treated is for example characterized by a sulfur content greater than 50 ppm and an olefin content greater than 10% for which it is desired to transform at least 70% sulfur to H ₂ S. the gas which has generally lower boiling temperatures than 250 ° C can be be treated by the device of the present invention, or undergo a pretreatment consisting of a selective hydrogenation step and fractionation . These pretreatments are described in detail in the application EP 1 077 247 . In this case, advantageously only the fraction C ₆₊ (that is to say containing the hydrocarbons whose total carbon number is greater than or equal to 6) of the gasoline can be treated by the process according to the present invention.

Gasoline mixed with hydrogen is heated by an exchanger train and / or an oven. The mixture brought to the desired temperature and pressure is usually in the vapor phase. It is sent to a first reactor (stage A) containing a hydrodesulphurization catalyst as described above implemented in a fixed bed. The effluent from this reactor contains hydrocarbons and sulfur compounds which have not reacted, paraffins resulting from the hydrogenation of olefins, H ₂ S resulting from the decomposition of sulfur compounds and recombinant mercaptans from H ₂ S addition reactions on olefins. This effluent is generally heated in an exchange train and / or an oven so that its temperature is increased by at least 10 ° C. and is injected into a second reactor (step B) containing a low-active hydrodesulfurization catalyst such as described above implemented in fixed bed. The effluent from this reactor consists of hydrocarbons and a reduced amount of sulfur compounds which have not reacted during step A, paraffins resulting from the hydrogenation of olefins, the H ₂ S from decomposition of sulfur compounds and a decreased amount of recombinant mercaptans from H ₂ S addition reactions to olefins.

The sequence of steps A and B allows, with respect to step A alone, to minimize the loss of olefins by hydrogenation, for a given desulfurization rate. The following examples illustrate the advantages of the process in one or two steps as just described. In these examples (as well as in the foregoing description), the contents of sulfur or of sulfur compounds are given in ppm by weight.

Example 1 Preparation of Catalysts

The catalysts are prepared according to the same method. The synthesis protocol consists in carrying out a dry impregnation of a solution of ammonium heptamolybdate and cobalt nitrate, the volume of the aqueous solution containing the metal precursors being equal to the volume of water recovery (ERV). corresponding to the mass of support to be impregnated.

The concentrations of the precursors in the solution are adjusted so as to deposit on the support the weight contents of desired metal oxides. The solid is then allowed to mature at room temperature for 12 hours and then dried at 120 ° C for 12 hours. Finally, the solid is calcined at 500 ° C. for two hours under an air flow (11 / h / g). The alumina supports used are industrial supports supplied by Axens, the characteristics of which are given in Table 1 below. <b> Table 1: </ b> Characteristics of industrial aluminums. Support Formatting S _BET (m ² / g) * V _p (Hg) ** cc / g α Balls 1.4 - 2.8 mm 140 1.10 β Balls 1.4 - 2.8 mm 80 1.09 γ Balls 1.4 - 2.8 mm 32 1.06 δ Balls 1.4 - 2.8 mm 210 0.64 * Specific surface area measured by nitrogen adsorption (ASTM D3663)
** total pore volume of Hg intrusion

Various CoMo type catalysts have been prepared on these supports. It can be seen (Table 2) that these catalysts are essentially distinguished by their textural properties for catalysts A, B, C, D and by their active phase content for catalysts E and F. <b> Table 2: </ b> characteristics of CoMo catalysts. Catalyst Support CoO% wt MoO ₃ % wt V (Hg) * cc / g Median pore diameter ** / nm A (compliant) α 3.5 10.0 0.99 22 B (compliant) β 3.5 9.2 0.87 54 C (compliant) γ 3.6 9.8 0.85 142 D (non-compliant) δ 3.8 10.7 0.60 12 E (non-compliant) δ 1.1 3.2 0.62 11 F (compliant) β 1.0 3.1 0.90 53 * total pore volume of Hg intrusion
** pore diameter corresponding to an intrusion volume of V _P (Hg) / 2

The sulphidation protocol of the catalyst is identical for each catalytic test. The catalyst, in its calcined form (oxide), is loaded into the catalytic test unit and then sulphurated by a synthetic filler (4% S in the form of DMDS in n-heptane). Sulfurization conditions are: LHSV = 2 h ^-1 (volume of feed / volume catalyst / hr), P = 2 MPa relative, H ₂ / feed = 300 (Nl / l) _bearing T = 350 ° C (4h, rise in T at 20 ° C / hour).

The sulfur content (in ppm) is evaluated in the feed and in the recipes (after removal of dissolved H ₂ S) according to the iso 14596 method; which allows the calculation of the rate of desulfurization of gasoline according to the formula: $$\mathrm{HDS}\left(\%\right)=\left(\mathrm{sulfur\; of\; the\; load\; in\; ppm}-\mathrm{sulfur\; of\; the\; recipe\; in\; ppm}\right)/\left(\mathrm{sulfur\; of\; the\; load\; in\; ppm}\right)*100$$

The weight content of olefins is evaluated in the feedstock and in the recipe by gas chromatography; this makes it possible to evaluate the degree of hydrogenation of the olefins of the gasoline according to the formula: $$\mathrm{HDO}\left(\%\right)=\left(\%\mathrm{Olefin\; weights\; charge}-\%\mathrm{olefin\; weights\; recipe}\right)/\left(\%\mathrm{Olefin\; weights\; charge}\right)*100$$

The total mercaptan content is measured in potentiometric recipes by the ASTM D3227 method after separation of H ₂ S.

Example 2: Evaluation of the Performance of Catalysts A and D

In this example the performances of catalysts A (compliant) and D (non-compliant) are compared in selective HDS of a sulfur-containing FCC gasoline whose characteristics are given in Table 3 below. <b> Table 3: </ b> Characteristics of FCC No. 1 Gasoline. Total sulfur (ppm): 970 Olefins (% by weight): 35.7 Aromatic (% weight): 27.6 ASTM distillation: PI: 37 ° C Mp 215 ° C

The test conditions are as follows: P = 2.7 relative MPa, VVH = 4 h ^-1 , H ₂ / load = 360 normal liters per liter (nl / l), T = 250 - 280 ° C. Each operating condition is maintained for the time necessary for the stabilization of the catalyst, both in hydrogenating activity and in desulfurizing activity (typically 24 to 48 hours). The results obtained on the catalysts A and D are shown in Table 4 below. <b> Table 4: </ b> Performance of Catalysts A and D for the Desulphurization of FCC No. 1 Gasoline. Catalyst A Catalyst D T (° C) 250 260 270 250 260 S _total 160 130 90 130 65 HDS /% 83.5 86.6 90.7 86.6 93.3 % olefins 26.7 26.1 25.5 23.0 21.1 HDO /% 25.2 26.9 28.6 35.6 40.9

It is observed that for comparable desulphurization levels (HDS), catalyst A has lower hydrogenation rates of olefins (HDO) than catalyst D. Catalyst A (compliant) is therefore more selective than catalyst D (no compliant)

Example 3: Evaluation of the Performance of Catalysts A and B.

In this example the catalysts A (compliant) and B (compliant) are evaluated on FCC No. 2 gasoline with lower sulfur content than FCC No. 1 gasoline, the characteristics of which are given in Table 5 below. . <b> Table 5: </ b> Characteristics of FCC No. 2 Gasoline. Total sulfur (ppm): 450 Olefins (% by weight): 33.5 Aromatic (% weight): 28.2 ASTM distillation: PI: -5 ° C PF: 252 ° C

The test conditions are as follows: P = 1.5 relative MPa, VVH = 5 h ^-1 , H ₂ / load = 300 Nl / l, T = 270-280 ° C. Each operating condition is maintained for the time necessary for the stabilization of the catalyst, both in hydrogenating activity and in desulfurizing activity (typically 24 to 48 hours). The results obtained on the catalysts A and B are shown in Table 6 below. <b> Table 6: </ b> Performance of Catalysts A and B for the Desulphurization of FCC No. 2 Gasoline. Catalyst A Catalyst B T (° C) 270 280 270 280 S _total 96 46 92 54 HDS /% 78.7 89.8 79.5 88.0 % olefins 29.7 26.3 30.1 27.5 HDO /% 11.3 21.5 10.1 17.9

For similar desulphurization rates (HDS), catalyst B has a lower hydrogenation activity (HDO) than catalyst A. Catalyst B (compliant) is therefore also more selective than catalyst D (non-compliant).

Example 4: Evaluation of the Performance of Catalysts A and C.

In this example, Catalyst A and Catalyst C are compared in desulphurization with a depentanized and highly highly sulfurized FCC # 3 gasoline whose characteristics are given in Table 7. <b> Table 7: </ b> Characteristics of FCC No. 3 Gasoline. Total sulfur (ppm): 2450 Olefins (% by weight): 32.1 Aromatic (% weight): 36.2 ASTM distillation: PI: 39 ° C PF: 240 ° C

The test conditions are as follows: P = 1.5 relative MPa, VVH = 4 h ^-1 , H ₂ / load = 300 Nl / l, T = 290-310 ° C. Each operating condition is maintained for the time necessary for the stabilization of the catalyst, both in hydrogenating activity and in desulfurizing activity (typically 24 to 96 hours). The results obtained on the catalysts A and C are shown in Table 8 below. <b> Table 8: </ b> Performance of Catalysts A and C for the Desulphurization of FCC No. 3 Gasoline Catalyst A Catalyst C T (° C) 290 310 290 310 Stotal 420 115 645 305 HDS /% 82.9 95.3 73.7 87.6 % olefins 25.0 19.7 27.1 23.2 HDO /% 22.1 38.6 15.6 27.7

The evolution of the degree of hydrogenation of the olefins as a function of the level of desulfurization shows that the two catalysts have comparable selectivities. Catalyst C is therefore no more selective than catalyst A. On the other hand, catalyst C is less active than catalyst A in hydrodesulfurization, which can potentially constitute a handicap in terms of the service life of this type of catalyst. an industrial unit. In terms of selectivity, catalyst C nevertheless remains higher than catalyst D (see Example 2, Table 4).

Example 5: Preparation of a Catalyst G partially deactivated

A 100 ml sample of catalyst B is subjected to accelerated deactivation on a pilot unit under the following conditions: the catalyst is operated at 300 ° C. under a mixture consisting of the gasoline 4 described in Example 6 and hydrogen injected up to 100 normal liters of hydrogen per liter of gasoline, with a gasoline flow rate of 400 ml / h and under a total pressure of 1 MPa relative. After 800 hours, the reactor is stripped at 120 ° C under nitrogen to remove the adsorbed hydrocarbons. The catalyst thus deactivated is called catalyst G.

Example 6 Evaluation of the Catalytic Activity of the Different Catalysts

The activity of the catalysts B, D, E, F, G is evaluated by a hydrodesulfurization test of a mixture of model molecules carried out in a stirred autoclave reactor of 500 ml. Typically between 2 g and 6 g of catalyst are sulfurized at atmospheric pressure in a sulfurization bench under H ₂ S / H ₂ mixture consisting of 15% by volume of H ₂ S at 1 l / h / g of catalyst and at 400 ° C. for two hours.

The model charge used for the activity test has the following composition: 1000 ppm of sulfur in thiophene form, 10% by weight of olefins in the form of 2,3-dimethyl-butene-2 in n-heptane.

This reaction mixture was chosen because considered representative of a catalytic cracking gasoline. The total system pressure is then adjusted and maintained at 3.5 relative MPa by hydrogen supply and the temperature is adjusted to 250 ° C. At time t = 0, the catalyst is contacted with the reaction mixture. Periodic sampling of samples makes it possible to follow the evolution of the composition of the solution over time by gas chromatographic analysis. The duration of the test is chosen so as to obtain final thiophene conversion levels of between 50 and 90%.

avec :

• A : activité du catalyseur, en min^-1 cm³ _catalyseur ^-1 (centimètre cube ^-1)
• k constante de vitesse pour la conversion du thiophène, en min^-1
• m_catalyseur : masse de catalyseur utilisée en g (avant sulfuration)
• DRT : densité de remplissage tassée du catalyseur, en cm³/g

The activity of the catalyst is defined with respect to the rate constant for conversion of normalized thiophene per volume of catalyst. The speed constant is calculated considering an order 1 for the reaction, $$\mathrm{AT}=\mathrm{k}/\left({\mathrm{m}}_{\mathrm{catalyst}}\times {\mathrm{DRT}}_{\mathrm{catalyst}}\right)$$

with:

• A: catalyst activity, in min ^-1 cm ³ _catalyst ^-1 (cubic centimeter ^-1 )
• k rate constant for thiophene conversion, in min ^-1
• m _catalyst : mass of catalyst used in g (before sulphurization)
• DRT: packed packing density of catalyst, in cm ³ / g

The relative activities of the catalysts B, D, E and F thus obtained are reported in Table 9 below, <b> Table 9: </ b> relative activities of catalysts B, D, E, F and G, Catalyst B Catalyst D Catalyst E Catalyst F Catalyst G Relative activity 100 * 120 42 31 45 * base

Example 7 Evaluation of the Performance of Catalysts B, D, E, F and G in a Sequence

Gasoline No. 4 described in Table 10 is used to study the performance of catalyst chains. This species is from an FCC unit and has been depentanized. <b> Table 10: </ b> Characteristics of FCC # 4 Gasoline Total sulfur (ppm): 380 Olefins (% by weight): 27.8 Olefins (% by weight): 32.1 Aromatic (% weight): 33.9 ASTM distillation: PI: 55 ° C PF: 219 ° C

The chaining tests are carried out in a pilot unit equipped with two reactors in series, each loaded with 100 ml of catalyst.

The performance of different catalyst sequences has been evaluated to illustrate the present invention. For each catalyst, the conventional sulfurization procedure described above was carried out, the same procedure for all the catalysts.

The basic operating conditions used for all the tests are as follows, namely a pressure equal to 1.8 MPa relative and a hydrogen-to-charge ratio of 400 normal liters per liter.

The temperatures were adjusted to reach a target of sulfur content between 10 ppm and 15 ppm. The following table 11 summarizes the performances of the various sequences evaluated. <b> Table 11: </ b> performance of catalysts alone or in sequence for the desulphurization of FCC No. 4 gasoline No. Test 1 2 3 4 5 6 7 catalysts B D D + E B + E B + F B + G B + D Temp. R1 28 27 275 280 280 280 280 Temp. R2 - - 300 300 300 300 275 VVH R1 (h ^-1 ) 4 4 8 8 8 8 8 VVH R2 (h ^-1 ) - 8 8 8 8 8 Overall VVH (h ^-1 ) 4 4 4 4 4 4 4 S effluent, ppm 12 13 14 13 15 12 13 Mercaptans, ppm 9 10 7 8 8 7 10 HDO% 28 32 24.5 21 20.1 21.4 30.6

The two reactors placed in series are named respectively reactor 1 and reactor 2. The volume of catalyst in each reactor is 100 ml.

Tests 1 and 2 were carried out on catalysts B and D alone. Catalyst D is not in accordance with the invention. The loss of olefins in run 1 is lower than the loss of olefins in run 2 due to the difference in selectivity between catalysts B and D.

The implementation of catalysts E, F or G in sequence with the catalysts B or D (tests 4, 5 in accordance with the invention) makes it possible to improve the overall selectivity. Indeed, for similar sulfur contents in the recipes, between 12 and 15 ppm, the losses of olefins measured by the HDO level are decreased compared to the tests 1 and 2, carried out on a single catalyst. In addition, observes that the best results have been obtained for the sequences 5 and 6 for which the catalysts used in the two steps are in accordance with the invention.

Test No. 7 is made from a sequence not according to the invention for which reactor 2 is charged with a more active catalyst than that charged to reactor 1. Compared to tests 3 to 6, , a loss of olefins and a higher residual mercaptan content, for equivalent sulfur content in effluents.

The comparison between the preceding tests also demonstrates the decrease in the amount of mercaptans in the product obtained by virtue of the practice of the invention.

Claims (14)

1. A process for hydrodesulphurizing a gasoline, comprising employing a catalyst comprising a support and an active phase comprising nickel or cobalt and in which the mean pore diameter of said catalyst is more than 20 nanometres,
   said process comprising at least two successive hydrodesulphurization steps in which said catalyst is employed in at least one of said steps,
   the activity of the catalyst of a step n+1 is in the range 1 percent to 90% of the activity of the catalyst of step n,
   the catalyst of step n+1 has a lower metals content than that of the catalyst of step n, and the successive steps being carried out without intermediate degassing
2. A process according to claim 1, in which the mean pore diameter is in the range 20 to 100 nm.
3. A process according to claim 1 or claim 2, in which the catalyst contains at least one group VI metal.
4. A process according to claim 3, in which the surface density of the group VI metal is in the range 2 x 10^-4 to 40 x 10^-4 grams of the oxide of said metal per m² of support.
5. A process according to one of the preceding claims, in which the support is selected from the group constituted by aluminas, silica, silica aluminas or oxides of titanium or magnesium, used alone or as a mixture with alumina or silica alumina.
6. A process according to one of the preceding claims, in which the support is at least in part constituted by an alumina.
7. A process according to one of the preceding claims, in which the specific surface area of the support is less than 200 m²/g,
8. A process according to claim 1, in which the reaction temperature of the successive steps is in the range from 150°C to 450°C, the pressure is in the range from 0.4 to 3 MPa relative, and the volume ratio of hydrogen to hydrocarbons, H₂/HC, is in the range from 50 N1/1 to 800 Nl/l.
9. A process according to claim 8 in which, for each step, the pressure is less than 2 MPa relative and the temperature is more than 250°C.
10. A process according to claim 1, in which the activity of a step is in the range 1 percent to 50 percent of the activity of the catalyst in the preceding step.
11. A process according to claim 1, in which the reaction temperature in step n+1 is higher than that in step n.
12. A process according to claims 1, in which the catalyst of step n+1 is the catalyst of step n which has undergone partial deactivation.
13. A process according to claim 12, in which the catalyst is deactivated by bringing the catalyst into contact with a feed containing a hydrocarbon fraction comprising olefins at a temperature of at least 250 °C.
14. A process according to claim 13, in which the catalyst of step n is recycled to step n+1 when its activity has reduced by at least 10 percent.