Patent Application
20250051661
The present invention relates to a process for the production of gasoline having a low content of sulfur and of mercaptans.
The production of gasolines meeting new environmental standards requires that their sulfur content be significantly decreased.
It is furthermore known that conversion gasolines, and more particularly those originating from catalytic cracking, which can represent from 30% to 50% of the gasoline pool, have high contents of monoolefins and of sulfur.
The sulfur present in gasolines is for this reason attributable, to close to 90%, to the gasolines resulting from catalytic cracking processes, which will be called FCC (Fluid Catalytic Cracking) gasolines subsequently. FCC gasolines thus constitute the preferred feedstock for the process of the present invention.
Among the possible routes for producing fuels having a low sulfur content, that which has been very widely adopted consists in specifically treating sulfur-rich gasoline bases by catalytic hydrodesulfurization processes in the presence of hydrogen. Conventional processes desulfurize gasolines in a nonselective manner by hydrogenating a large part of the monoolefins, which causes a high loss in octane number and a high consumption of hydrogen. The most recent processes, such as the Prime G+ (trademark) process, make it possible to desulfurize cracked gasolines rich in olefins, while limiting the hydrogenation of the monoolefins and consequently the loss of octane and the high hydrogen consumption which results therefrom. Such processes are, for example, described in the patent applications EP 1 077 247 and EP 1 174 485.
The residual sulfur compounds generally present in desulfurized gasoline can be separated into two distinct families: the unconverted refractory sulfur compounds present in the feedstock, on the one hand, and the sulfur compounds formed in the reactor by secondary “recombination” reactions. Among the latter family of sulfur compounds, the predominant compounds are the mercaptans resulting from the addition of H2S formed in the reactor to the monoolefins present in the feedstock.
Mercaptans, of chemical formula R—SH, where R is an alkyl group, are also called recombinant mercaptans. Their formation or their decomposition obeys the thermodynamic equilibrium of the reaction between monoolefins and hydrogen sulfide to form recombinant mercaptans. An example is illustrated according to the following reaction:
The sulfur contained in the recombinant mercaptans generally represents between 20% and 80% by weight of the residual sulfur in desulfurized gasolines.
The formation of recombinant mercaptans is in particular described in the patent U.S. Pat. No. 6,231,754 and the patent application WO01/40409, which teach various combinations of operating conditions and of catalysts making it possible to limit the formation of recombinant mercaptans.
Other solutions to the problem of the formation of recombinant mercaptans are based on a treatment of partially desulfurized gasolines in order to extract therefrom said recombinant mercaptans. Some of these solutions are described in the patent applications WO02/28988 or WO01/79391.
Still other solutions are described in the literature for desulfurizing cracked gasolines using a combination of steps of hydrodesulfurization and of removal of the recombinant mercaptans by reaction to give thioethers or disulfides (also called sweetening) (see, for example, U.S. Pat. Nos. 7,799,210, 6,960,291, US2007/114156, EP 2 861 094).
Finally, patent application US2006/278567 discloses a process for hydrodesulfurization of a cracked gasoline in two steps comprising intermediate separation steps making it possible notably to limit the formation of mercaptans.
A subject of the present invention is a process for treating a gasoline containing sulfur compounds, olefins and diolefins, the process comprising at least the following steps:
The applicant has identified, surprisingly, that in order to maintain a mercaptan and sulfur content in the hydrodesulfurized gasoline of less than 10 ppm by weight while limiting the loss of octane, the pressure in the hydrodesulfurization section (step a) of the process) must be between 1.5 and 3 MPa and the separation drum pressure must be within a specific range between 1.0 and 2.0 MPa. In fact, if the pressure of the separation drum is less than 1.0 MPa, the purity of the hydrogen contained in the gaseous recycle fraction will decrease, thus leading to a decrease in the ratio between the hydrogen flow rate and the flow rate of the feedstock to be treated. Thus, in order to maintain a mercaptan and sulfur content of 10 ppm by weight, it is necessary to increase the temperature of the hydrodesulfurization steps, which has the effect of increasing the hydrogenation of the olefins of the gasoline cut, which will have the effect of increasing the loss of octane. Moreover, if the pressure of the separation drum is greater than 2.0 MPa, the purity of the hydrogen contained in the gaseous recycle fraction will certainly increase, but this leads to an increase in the pressure at the inlet of the hydrodesulfurization reactors and therefore, in addition to an excess consumption of hydrogen, will result in an increase in mercaptan content in hydrodesulfurized gasoline and an increase in loss of octane.
According to one or more embodiments, the pressure of the separation drum of step c) is between 1.2 and 1.8 MPa.
According to one or more embodiments, the catalyst of step a) comprises a group VIII metal content of between 0.1% and 10% by weight of oxide of the group VIII metal relative to the total weight of the catalyst, and a group VIB metal content of between 1% and 20% by weight of oxide of the group VIB metal relative to the total weight of the catalyst.
According to one or more embodiments, the catalyst of step a) comprises a molar ratio of group VIII metal to group VIB metal of the catalyst of between 0.1 and 0.8.
According to one or more embodiments, the catalyst of step a) comprises a specific surface area of between 5 and 400 m2/g.
According to one or more embodiments, the catalyst of step a) comprises alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight, with respect to the total weight of catalyst, of cobalt oxide, in CoO form, of between 0.1% and 10%, a content by weight, with respect to the total weight of catalyst, of molybdenum oxide, in MoO3 form, of between 1% and 20%, a cobalt/molybdenum molar ratio of between 0.1 and 0.8 and a content by weight, with respect to the total weight of catalyst, of phosphorus oxide in P2O5 form of between 0.3% and 10% when phosphorus is present, said catalyst having a specific surface area of between 50 and 250 m2/g.
According to one or more embodiments, the catalyst of step b) comprises a group VIII metal content of between 1% and 60% by weight of oxide of the group VIII metal relative to the total weight of the catalyst.
According to one or more embodiments, the catalyst of step b) comprises a specific surface area of between 5 and 400 m2/g.
According to one or more embodiments, the catalyst of step b) consists of alumina and of nickel, said catalyst containing a content by weight, with respect to the total weight of catalyst, of nickel oxide, in NiO form, of between 5% and 20%, said catalyst having a specific surface area of between 30 and 180 m2/g.
According to one or more embodiments, the temperature of step b) is higher than the temperature of step a).
According to one or more embodiments, the temperature of step b) is at least 5° C. higher than the temperature of step a).
According to one or more embodiments, before step a), a gasoline distillation step is carried out so as to fractionate said gasoline into at least two, light and heavy, gasoline cuts and the heavy gasoline cut is treated in steps a), b), c), d) and e).
According to one or more embodiments, before step a) and before any optional distillation step, the gasoline is brought into contact with hydrogen and a selective hydrogenation catalyst in order to selectively hydrogenate the diolefins contained in said gasoline to give olefins.
According to one or more embodiments, the gasoline is a catalytic cracking gasoline.
Subsequently, the groups of chemical elements are given according to the CAS classification (CRC Handbook of Chemistry and Physics, published by CRC Press, editor-in-chief D. R. Lide, 81st edition, 2000-2001). For example, group VIII according to the CAS classification corresponds to the metals of columns 8, 9 and 10 according to the new IUPAC classification.
The term “specific surface area” is understood to mean the BET specific surface area (SBET in m2/g) determined by nitrogen adsorption in accordance with the standard ASTM D 3663-78 established from the Brunauer-Emmett-Teller method described in the journal “The Journal of the American Chemical Society”, 1938, 60, 309.
Total pore volume of the catalyst or of the support used for the preparation of the catalyst is understood to mean the volume measured by mercury porosimetry intrusion according to the standard ASTM D4284 at a maximum pressure of 4000 bar (400 MPa), using a surface tension of 484 dynes/cm and a contact angle of 140°, for example with a Microméritics® instrument, model Autopore III.
The wetting angle used was taken as equal to 140° following the recommendations of the publication “Techniques de l′ingénieur, traité analyse et caractérisation” [Techniques of the Engineer, Analysis and Characterization Treatise], pages 1050-1055, written by Jean Charpin and Bernard Rasneur. In order to obtain better accuracy, the value of the total pore volume corresponds to the value of the total pore volume measured by mercury intrusion porosimetry measured on the sample minus the value of the total pore volume measured by mercury intrusion porosimetry measured on the same sample for a pressure corresponding to 30 psi (approximately 0.2 MPa).
The contents of group VIII elements, group VIB elements and phosphorus are measured by X-ray fluorescence.
The process according to the invention makes it possible to treat any type of gasoline cut containing sulfur-comprising compounds and olefins, such as, for example, a cut resulting from a coking, visbreaking, steam cracking or catalytic cracking (FCC, Fluid Catalytic Cracking) unit. This gasoline can optionally be composed of a significant fraction of gasoline originating from other production processes, such as atmospheric distillation (gasoline resulting from a direct distillation (or straight run gasoline)), or from conversion processes (coking or steam cracking gasoline). Said feedstock preferably consists of a gasoline cut resulting from a catalytic cracking unit.
The feedstock is a gasoline cut containing sulfur compounds and olefins, the boiling point range of which typically extends from the boiling points of the hydrocarbons having 2 or 3 carbon atoms (C2 or C3) up to 260° C., preferably from the boiling points of the hydrocarbons having 2 or 3 carbon atoms (C2 or C3) up to 220° C., more preferably from the boiling points of the hydrocarbons having 5 carbon atoms up to 220° C. The process according to the invention can also treat feedstocks having lower end points than those mentioned above, such as, for example, a C5-180° C. cut.
The sulfur content of the gasoline cuts produced by catalytic cracking (FCC) depends on the sulfur content of the feedstock treated by the FCC, on the presence or absence of a pretreatment of the FCC feedstock, and also on the end point of the cut. Generally, the sulfur contents of the whole of a gasoline cut, notably those originating from FCC, are greater than 100 ppm by weight and most of the time greater than 500 ppm by weight. For gasolines having end points of greater than 200° C., the sulfur contents are often greater than 1000 ppm by weight; they can even, in certain cases, reach values of the order of 4000 to 5000 ppm by weight.
The feedstock treated by the process according to the invention can be a feedstock containing sulfur compounds in a content of greater than 1000 ppm by weight of sulfur and possibly greater than 1500 ppm.
In addition, the gasolines resulting from catalytic cracking (FCC) units contain, on average, between 0.5% and 5% by weight of diolefins, between 20% and 50% by weight of olefins and between 10 ppm and 0.5% by weight of sulfur, generally including less than 300 ppm of mercaptans.
The hydrodesulfurization step a) is implemented in order to reduce the sulfur content of the gasoline to be treated by converting the sulfur compounds into H2S, which is subsequently removed in step c).
The hydrodesulfurization step a) consists in bringing the gasoline to be treated into contact with hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization.
According to a preferred embodiment of the invention, step a) is implemented with the aim of carrying out a hydrodesulfurization selectively, that is to say with a degree of hydrogenation of the monoolefins of less than 80%, preferably of less than 70% and very preferably of less than 60%.
The temperature is generally between 21° and 320° C. and preferably between 22° and 290° C. The temperature employed must be sufficient to maintain the gasoline to be treated in the vapor phase in the reactor. In the case where the hydrodesulfurization step a) is carried out in several reactors in series, the temperature of each reactor is generally greater by at least 5° C., preferably by at least 10° C. and very preferably by at least 30° C. than the temperature of the reactor which precedes it.
The operating pressure of this step is generally between 1.5 and 3 MPa.
The amount of catalyst employed in each reactor is generally such that the ratio of the flow rate of gasoline to be treated, expressed in m3 per hour under standard conditions, per m3 of catalyst (also called space velocity) is between 1 and 10 h−1 and preferably between 2 and 8 h 1.
The hydrogen flow rate is generally such that the ratio of the hydrogen flow rate, expressed in normal m3 per hour (Nm3/h), to the flow rate of feedstock to be treated, expressed in m3 per hour under standard conditions (15° C., 0.1 MPa), is between 100 and 600 Nm3/m3, preferably between 200 and 500 Nm3/m3. Normal m3 is understood to mean the amount of gas in a volume of 1 m3 at 0° C. and 0.1 MPa.
The hydrogen required for this step can be fresh hydrogen or recycled hydrogen, preferably freed from H2S, or a mixture of fresh hydrogen and of recycled hydrogen. Preferably, fresh hydrogen will be used.
The degree of desulfurization of step a), which depends on the sulfur content of the feedstock to be treated, is generally greater than 50% and preferably greater than 70%, so that the product resulting from step a) contains less than 100 ppm by weight of sulfur and preferably less than 50 ppm by weight of sulfur.
The catalyst used in step a) must exhibit a good selectivity with regard to the hydrodesulfurization reactions, in comparison with the reaction for the hydrogenation of olefins.
The hydrodesulfurization catalyst of step a) comprises an oxide support and an active phase comprising a group VIB metal and a group VIII metal and optionally phosphorus and/or an organic compound as described below.
The group VIB metal present in the active phase of the catalyst is preferentially chosen from molybdenum and tungsten. The group VIII metal present in the active phase of the catalyst is preferentially chosen from cobalt, nickel and the mixture of these two elements. The active phase of the catalyst is preferably chosen from the group formed by the combination of the elements nickel-molybdenum, cobalt-molybdenum and nickel-cobalt-molybdenum and very preferably the active phase consists of cobalt and molybdenum.
The group VIII metal content is between 0.1% and 10% by weight of oxide of the group VIII metal relative to the total weight of the catalyst, preferably between 0.6% and 8% by weight, preferably between 0.6% and 7% by weight, very preferably between 1% and 6% by weight. The content of group VIB metal is between 1% and 20% by weight of oxide of the group VIB metal, with respect to the total weight of the catalyst, preferably between 2% and 18% by weight, very preferably between 3% and 16% by weight.
The group VIII metal to group VIB metal molar ratio of the catalyst is generally between 0.1 and 0.8, preferably between 0.2 and 0.6.
Optionally, the catalyst can additionally exhibit a phosphorus content generally of between 0.3% and 10% by weight of P2O5, with respect to the total weight of catalyst, preferably between 0.3% and 5% by weight, very preferably between 0.5% and 3% by weight. For example, the phosphorus present in the catalyst is combined with the group VIB metal and optionally also with the group VIII metal in the form of heteropolyanions.
Furthermore, the phosphorus/(group VIB metal) molar ratio is generally between 0.1 and 0.7, preferably between 0.2 and 0.6, when phosphorus is present.
Preferably, the catalyst is characterized by a specific surface area of between 5 and 400 m2/g, preferably between 10 and 250 m2/g and preferably between 50 and 250 m2/g. The specific surface area is determined in the present invention by the BET method according to the standard ASTM D3663, as described in the work by Rouquerol F., Rouquerol J. and Singh K., Adsorption by Powders & Porous Solids: Principle, Methodology and Applications”, Academic Press, 1999, for example by means of a Micromeritics™ brand Autopore III™ model machine.
The total pore volume of the catalyst is generally between 0.4 cm3/g and 1.3 cm3/g, preferably between 0.6 cm3/g and 1.1 cm3/g. The total pore volume is measured by mercury porosimetry according to the standard ASTM D4284 with a wetting angle of 140°, as described in the same work.
The tapped packing density (TPD) of the catalyst is generally between 0.4 and 0.8 g/ml, preferably between 0.4 and 0.7 g/ml. The TPD measurement consists in introducing the catalyst into a measuring cylinder, the volume of which has been determined beforehand, and then, by vibration, in tapping it until a constant volume is obtained. The bulk density of the tapped product is calculated by comparing the mass introduced and the volume occupied after tapping.
The catalyst can be in the form of cylindrical or multilobe (trilobe, quadrilobe, and the like) extrudates with a small diameter, or of spheres.
The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas and also titanium or magnesium oxides, used alone or as a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, the family of the transition aluminas and silica-aluminas; very preferably, the oxide support is constituted essentially of alumina, that is to say that it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, indeed even at least 90% by weight, of alumina. It preferably consists solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, that is to say which contains theta-, delta-, kappa- or alpha-phase aluminas, alone or as a mixture, and an amount of less than 20% of gamma-, chi- or eta-phase alumina.
The catalyst can also additionally comprise at least one organic compound containing oxygen and/or nitrogen and/or sulfur before sulfidation.
A very preferred embodiment of the invention corresponds to the use, for step a), of a catalyst comprising alumina and an active phase comprising cobalt, molybdenum and optionally phosphorus, said catalyst containing a content by weight, with respect to the total weight of catalyst, of cobalt oxide, in CoO form, of between 0.1% and 10%, a content by weight, with respect to the total weight of catalyst, of molybdenum oxide, in MoO3 form, of between 1% and 20%, a cobalt/molybdenum molar ratio of between 0.1 and 0.8 and a content by weight, with respect to the total weight of catalyst, of phosphorus oxide in P2O5 form of between 0.3% and 10% when phosphorus is present, said catalyst having a specific surface area of between 50 and 250 m2/g. According to one embodiment, the active phase consists of cobalt and molybdenum. According to another embodiment, the active phase consists of cobalt, molybdenum and phosphorus.
During the hydrodesulfurization step a), a large part of the sulfur compounds is converted into H2S. The remaining sulfur compounds are essentially refractory sulfur compounds and the recombinant mercaptans resulting from the addition of H2S formed in step a) to the monoolefins present in the feedstock.
The “finishing” hydrodesulfurization step is mainly implemented to reduce the content of the recombinant mercaptans. Preferably, step b) is carried out at a higher temperature than that of step a). Specifically, by using a higher temperature in this step compared to the temperature of step a), the formation of olefins and of H2S will be promoted by the thermodynamic equilibrium. Step b) also makes it possible to hydrodesulfurize the more refractory sulfur compounds.
The hydrodesulfurization step b) consists in bringing the effluent from step a) into contact optionally with an addition of hydrogen, in one or more hydrodesulfurization reactors, containing one or more catalysts suitable for carrying out the hydrodesulfurization.
The hydrodesulfurization step b) is carried out without significant hydrogenation of the olefins. The degree of hydrogenation of the olefins of the catalyst of the hydrodesulfurization step b) is generally less than 5% and more generally still less than 2%.
The temperature of this step is generally between 28° and 400° C., more preferably between 29° and 380° C. and very preferably between 30° and 360° C. The temperature of this step b) is generally greater by at least 5° C., preferably by at least 10° C. and very preferably by at least 30° C. than the temperature of step a).
The operating pressure of this step is generally between 1.0 and 3 MPa and preferably between 1.5 and 3 MPa.
The amount of catalyst employed in each reactor is generally such that the ratio of the flow rate of gasoline to be treated, expressed in m3 per hour under standard conditions, per m3 of catalyst (also called space velocity) is between 1 and 10 h−1 and preferably between 2 and 8 h−1.
Preferably, the hydrogen flow rate is subject and equal to the amount injected in step a) decreased by the hydrogen consumed in step a). The hydrogen flow rate is generally such that the ratio of the hydrogen flow rate, expressed in normal m3 per hour (Nm3/h), to the flow rate of feedstock to be treated, expressed in m3 per hour at standard conditions (15° C., 0.1 MPa), is between 100 and 600 Nm3/m3, preferably between 200 and 500 Nm3/m3.
The degree of desulfurization in step b), which depends on the sulfur content of the feedstock to be treated, is generally greater than 50% and preferably greater than 70%, so that the product resulting from step b) contains less than 60 ppm by weight of sulfur, preferably less than 40 ppm by weight of sulfur, and very preferably less than 20 ppm by weight of sulfur.
The hydrodesulfurization steps a) and b) may be carried out either in a single reactor containing both catalysts or in at least two different reactors. When steps a) and b) are carried out using two reactors, these two reactors are placed in series with the second reactor treating all of the effluent exiting from the first reactor (without separation of the liquid and of the gas between the first and second reactor).
The catalyst of step b) is different in nature and/or in composition from that used in step a). The catalyst of step b) is in particular a very selective hydrodesulfurization catalyst: it makes it possible to hydrodesulfurize without hydrogenating the olefins and thus to maintain the octane number.
The catalyst which may be suitable for this step b) of the process according to the invention, without this list being limiting, is a catalyst comprising an oxide support and an active phase consisting of at least one group VIII metal and preferably chosen from the group formed by nickel, cobalt and iron. These metals can be used alone or in combination. Preferably, the active phase consists of a group VIII metal, preferably nickel. Particularly preferably, the active phase consists of nickel.
The content of group VIII metal is between 1% and 60% by weight of oxide of the group VIII metal, with respect to the total weight of the catalyst, preferably between 5% and 30% by weight, very preferably between 5% and 20% by weight.
Preferably, the catalyst is characterized by a specific surface area of between 5 and 400 m2/g, preferably of between 10 and 250 m2/g, preferably of between 20 and 200 m2/g, very preferably of between 30 and 180 m2/g. The specific surface area is determined in the present invention by the BET method according to the standard ASTM D3663, as described in the work by Rouquerol F., Rouquerol J. and Singh K., Adsorption by Powders & Porous Solids: Principle, Methodology and Applications”, Academic Press, 1999, for example by means of a Micromeritics™ brand Autopore III™ model machine.
The pore volume of the catalyst is generally between 0.4 cm3/g and 1.3 cm3/g, preferably between 0.6 cm3/g and 1.1 cm3/g. The total pore volume is measured by mercury porosimetry according to the standard ASTM D4284 with a wetting angle of 140°, as described in the same work.
The tapped packing density (TPD) of the catalyst is generally between 0.4 and 0.8 g/ml, preferably between 0.4 and 0.7 g/ml.
The TPD measurement consists in introducing the catalyst into a measuring cylinder, the volume of which has been determined beforehand, and then, by vibration, in tapping it until a constant volume is obtained. The bulk density of the tapped product is calculated by comparing the mass introduced and the volume occupied after tapping.
The catalyst can be in the form of cylindrical or multilobe (trilobe, quadrilobe, and the like) extrudates with a small diameter, or of spheres.
The oxide support of the catalyst is usually a porous solid chosen from the group consisting of: aluminas, silica, silica-aluminas and also titanium or magnesium oxides, used alone or as a mixture with alumina or silica-alumina. It is preferably chosen from the group consisting of silica, the family of the transition aluminas and silica-aluminas; very preferably, the oxide support is constituted essentially of alumina, that is to say that it comprises at least 51% by weight, preferably at least 60% by weight, very preferably at least 80% by weight, indeed even at least 90% by weight, of alumina. It preferably consists solely of alumina. Preferably, the oxide support of the catalyst is a “high temperature” alumina, that is to say which contains theta-, delta-, kappa- or alpha-phase aluminas, alone or as a mixture, and an amount of less than 20% of gamma-, chi- or eta-phase alumina.
A very preferred embodiment of the invention corresponds to the use, for step b), of a catalyst consisting of alumina and of nickel, said catalyst containing a content by weight, with respect to the total weight of catalyst, of nickel oxide, in NiO form, of between 5% and 20%, said catalyst having a specific surface area of between 30 and 180 m2/g.
The catalyst of the hydrodesulfurization step b) is characterized by a hydrodesulfurization catalytic activity generally of between 1% and 90%, preferentially of between 1% and 70% and very preferably of between 1% and 50% of the catalytic activity of the catalyst of the hydrodesulfurization step a).
This step is carried out in order to separate the excess hydrogen and also the H2S formed during steps a) and b) and to extract the desulfurized gasoline meeting the sulfur and recombination mercaptan specifications while reducing the losses of C5+ hydrocarbon compounds (octane).
According to step c) of the process according to the invention, the effluent of step b) is fractionated so as to produce a gas phase comprising hydrogen from steps a) and b), H2S formed during steps a) and b) and C1 to C4 light compounds, and desulfurized gasoline.
More particularly, after the hydrodesulfurization step b), the gasoline is advantageously cooled to a temperature generally of less than 80° C. and preferably less than 60° C., in order to condense the hydrocarbons. The gas and liquid phases are subsequently separated in a separation drum. The temperature of the separation drum is generally between 2° and 80° C., preferably between 25 and 65° C. The pressure of the separation drum is fixed between 1.0 and 2.0 MPa, preferably between 1.2 and 1.8 MPa.
Step c) is preferably carried out in order for the sulfur in the form of H2S remaining in the effluent from step b) to represent less than 30%, preferably less than 20% and more preferably less than 10% of the total sulfur present in the treated hydrocarbon fraction.
The liquid fraction, which contains the desulfurized gasoline and also a fraction of the H2S dissolved, is sent to a stabilization column or debutanizer. This column separates a top cut, consisting essentially of residual H2S and of hydrocarbon compounds having a boiling point less than or equal to that of butane (C4-), and a bottom cut freed from H2S, which corresponds to the stabilized desulfurized gasoline, containing the compounds having a boiling point greater than that of butane.
The use of a separation drum operating advantageously at low temperature to remove H2, H2S and the light gases, followed by a stabilization column to remove the dissolved H2S and to produce the stabilized desulfurized gasoline reduces the temperature at the bottom of the stabilization column compared to separation of H2 and H2S directly in a stabilization column.
The stabilization column generally operates at a pressure of between 0.1 and 2 MPa, preferably between 0.2 and 1 MPa.
The number of theoretical plates of this separation column is generally between 10 and 50 and preferably between 20 and 40.
The reflux ratio, expressed as being the ratio of the liquid flow in the column divided by the distillate flow rate expressed in kg/h, is generally less than unity and preferably less than 0.5.
The desulfurized and stabilized heavy gasoline produced by the process according to the invention is advantageously used as a base for formulating a gasoline fuel.
According to the invention, at least one portion of the gaseous fraction separated by the separation drum in step c) is recycled to steps a) and/or b), preferably to step b).
Advantageously, said gaseous fraction is sent beforehand to an amine absorber or a washing column operating at low pressure (typically at 1.5 MPa), in order to remove at least one portion of the H2S.
The preparation of the catalysts of steps a) and b) is known and generally comprises a step of impregnation of the metals from group VIII and from group VIB, when it is present, and optionally of phosphorus and/or of the organic compound on the oxide support, followed by a drying operation and then by an optional calcination making it possible to obtain the active phase in their oxide forms. Before its use in a process for the hydrodesulfurization of a sulfur-containing olefinic gasoline cut, the catalysts are generally subjected to a sulfidation in order to form the active entity as described below.
The impregnation step can be carried out either by slurry impregnation, or by impregnation in excess, or by dry impregnation, or by any other means known to a person skilled in the art.
The impregnation solution is chosen so as to be able to dissolve the metal precursors in the desired concentrations.
Use may be made, by way of example, among the sources of molybdenum, of the oxides and hydroxides, molybdic acids and salts thereof, in particular the ammonium salts, such as ammonium molybdate, ammonium heptamolybdate, phosphomolybdic acid (H3PMo12O40), and salts thereof, and optionally silicomolybdic acid (H4SiMo12O40) and salts thereof. The sources of molybdenum can also be any heteropolycompound of Keggin, lacunary Keggin, substituted Keggin, Dawson, Anderson or Strandberg type, for example. Use is preferably made of molybdenum trioxide and the heteropolycompounds of Keggin, lacunary Keggin, substituted Keggin and Strandberg type.
The tungsten precursors which can be used are also well known to a person skilled in the art. For example, use may be made, among the sources of tungsten, of the oxides and hydroxides, tungstic acids and their salts, in particular the ammonium salts, such as ammonium tungstate or ammonium metatungstate, phosphotungstic acid and their salts, and optionally silicotungstic acid (H4SiW12O40) and its salts. The sources of tungsten can also be any heteropolycompound of Keggin, lacunary Keggin, substituted Keggin or Dawson type, for example. Use is preferably made of the oxides and the ammonium salts, such as ammonium metatungstate, or the heteropolyanions of Keggin, lacunary Keggin or substituted Keggin type.
The cobalt precursors which can be used are advantageously chosen from the oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example. Use is preferably made of cobalt hydroxide and cobalt carbonate.
The nickel precursors which can be used are advantageously chosen from the oxides, hydroxides, hydroxycarbonates, carbonates and nitrates, for example.
The preferred phosphorus precursor is orthophosphoric acid H3PO4 but its salts and esters, such as ammonium phosphates, are also suitable. The phosphorus can also be introduced at the same time as the element(s) from group VIB in the form of Keggin, lacunary Keggin, substituted Keggin or Strandberg-type heteropolyanions.
After the impregnation step, the catalyst is generally subjected to a drying step at a temperature of less than 200° C., advantageously of between 50° C. and 180° C., preferably between 70° C. and 150° C., very preferably between 75° C. and 130° C. The drying step is preferentially carried out under an inert atmosphere or under an oxygen-containing atmosphere. The drying step may be carried out by any technique known to a person skilled in the art. It is advantageously carried out at atmospheric pressure or at reduced pressure. Preferably, this step is carried out at atmospheric pressure. It is advantageously carried out in a traversed bed using hot air or any other hot gas. Preferably, when the drying is carried out in a fixed bed, the gas used is either air or an inert gas, such as argon or nitrogen. Very preferably, the drying is carried out in a traversed bed in the presence of nitrogen and/or of air. Preferably, the drying step has a duration of between 5 minutes and 15 hours, preferably between 30 minutes and 12 hours.
According to a variant of the invention, the catalyst has not undergone calcination during its preparation, that is to say that the impregnated catalytic precursor has not been subjected to a step of heat treatment at a temperature of greater than 200° C. under an inert atmosphere or under an oxygen-containing atmosphere, in the presence or absence of water.
According to another variant of the invention, which is preferred, the catalyst has undergone a calcination step during its preparation, that is to say that the impregnated catalytic precursor has been subjected to a step of heat treatment at a temperature of between 250° C. and 1000° C. and preferably between 200° C. and 750° C., for a period of time typically of between 15 minutes and 10 hours, under an inert atmosphere or under an oxygen-containing atmosphere, in the presence or absence of water.
Before bringing into contact with the feedstock to be treated in a process for the hydrodesulfurization of gasolines, the catalysts of the process according to the invention generally undergo a sulfidation step. The sulfidation is preferably carried out in a sulforeducing medium, that is to say in the presence of H2S and of hydrogen, in order to transform the metal oxides into sulfides, such as, for example, MoS2, Co9S8 or Ni3S2. The sulfidation is carried out by injecting, onto the catalyst, a stream containing H2S and hydrogen, or else a sulfur compound capable of decomposing to give H2S in the presence of the catalyst and of hydrogen. Polysulfides, such as dimethyl disulfide (DMDS), are H2S precursors commonly used to sulfide catalysts. The sulfur can also originate from the feedstock. The temperature is adjusted in order for the H2S to react with the metal oxides to form metal sulfides. This sulfidation can be carried out in situ or ex situ (inside or outside the reactor) of the reactor of the process according to the invention at temperatures of between 200° C. and 600° C. and more preferentially between 300° C. and 500° C.
The degree of sulfidation of the metals constituting the catalysts is at least equal to 60%, preferably at least equal to 80%. The sulfur content in the sulfided catalyst is measured by elemental analysis according to ASTM D5373. A metal is regarded as sulfided when the overall degree of sulfidation, defined by the molar ratio of the sulfur(S) present on the catalyst to said metal, is at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of the metal(s) under consideration. The overall degree of sulfurization is defined by the following equation:
wherein:
(S/metal)catalyst is the molar ratio of sulfur(S) to metal present on the catalyst
(S/metal)theoretical is the molar ratio of sulfur to metal corresponding to the complete sulfidation of the metal to give sulfide.
This theoretical molar ratio varies according to the metal under consideration:
When the catalyst comprises several metals, the molar ratio of S present on the catalyst to the combined metals also has to be at least equal to 60% of the theoretical molar ratio corresponding to the complete sulfidation of each metal to give sulfide, the calculation being carried out in proportion to the relative molar fractions of each metal.
Schemes which can be Employed within the Scope of the Invention
Different schemes can be employed in order to produce, at a lower cost, a desulfurized gasoline having a reduced content of mercaptans. The choice of the optimum scheme depends in fact on the characteristics of the gasolines to be treated and to be produced and also on the constraints specific to each refinery.
The schemes described below are given by way of illustration without limitation.
According to a first variant, a step of distillation of the gasoline to be treated is carried out in order to separate two cuts (or fractions), namely a light cut and a heavy cut, and the heavy cut is treated according to the process of the invention. The light cut generally has a boiling point range of less than 100° C. and the heavy cut a temperature range of greater than 65° C. This first variant has the advantage of not hydrotreating the light cut which is rich in olefins and generally comprises a low sulfur content, which makes it possible to limit the loss of octane through hydrogenation of the olefins.
According to a second variant, the gasoline to be treated is subjected, before the process according to the invention, to a preliminary step consisting of a selective hydrogenation of the diolefins present in the feedstock, as described in the patent application EP 1 077 247.
The gasoline to be treated is treated beforehand in the presence of hydrogen and of a selective hydrogenation catalyst so as to at least partially hydrogenate the diolefins and to carry out a reaction for increasing the molecular weight of a portion of the light mercaptan (RSH) compounds present in the feedstock to give thioethers, by reaction with olefins. To this end, the gasoline to be treated is sent to a selective hydrogenation catalytic reactor containing at least one fixed or moving bed of catalyst for the selective hydrogenation of the diolefins and for increasing the molecular weight of the light mercaptans. The reaction for the selective hydrogenation of the diolefins and for increasing the molecular weight of the light mercaptans is preferentially carried out on a sulfided catalyst comprising at least one element from group VIII and optionally at least one element from group VIB and an oxide support. The element from group VIII is preferably chosen from nickel and cobalt and in particular nickel. The element from group VIb, when it is present, is preferably chosen from molybdenum and tungsten and very preferably molybdenum.
The oxide support of the catalyst is preferably chosen from alumina, nickel aluminate, silica, silicon carbide or a mixture of these oxides. Use is preferably made of alumina and more preferably still of high-purity alumina. According to a preferred embodiment, the selective hydrogenation catalyst contains nickel at a content by weight of nickel oxide, in NiO form, of between 1% and 12%, and molybdenum at a content by weight of molybdenum oxide, in MoO3 form, of between 6% and 18% and a nickel/molybdenum molar ratio of between 0.3 and 2.5, the metals being deposited on a support consisting of alumina. The degree of sulfidation of the metals constituting the catalyst is preferably greater than 60%.
During the optional selective hydrogenation step, the gasoline is brought into contact with the catalyst at a temperature of between 5° and 250° C., preferably between 8° and 220° C. and more preferably still between 9° and 200° C., with a liquid space velocity of between 0.5 h−1 and 20 h−1, the unit of the liquid space velocity being the liter of feedstock per liter of catalyst and per hour (I/I/h). The pressure is between 0.4 and 5 MPa, preferably between 0.6 and 4 MPa and more preferably still between 1 and 3 MPa. The optional selective hydrogenation step is typically carried out with a ratio of the hydrogen flow rate, expressed in normal m3 per hour, to the flow rate of feedstock to be treated, expressed in m3 per hour at standard conditions, of between 2 and 100 Nm3/m3, preferably between 3 and 30 Nm3/m3.
After selective hydrogenation, the content of diolefins, determined via the maleic anhydride value (MAV), according to the UOP 326 method, is generally reduced to less than 6 mg maleic anhydride/g, indeed even less than 4 mg MA/g and more preferably less than 2 mg MA/g. In some cases, there may be obtained less than 1 mg MA/g.
The selectively hydrogenated gasoline is subsequently distilled into at least two cuts, a light cut and a heavy cut and optionally an intermediate cut. In the case of the fractionation into two cuts, the heavy cut is treated according to the process of the invention. In the case of the fractionation into three cuts, the intermediate and heavy cuts can be treated separately by the process according to the invention.
It should be noted that it is possible to envisage carrying out the steps of hydrogenation of the diolefins and of fractionation in two or three cuts simultaneously by means of a catalytic distillation column which includes a distillation column equipped with at least one catalytic bed.
Other characteristics and advantages of the invention will now become apparent on reading the description which will follow, given solely by way of illustration and without limitation, and with reference to the appended
With reference to
Then, the heavy gasoline cut is sent via the line 7 and hydrogen is sent via the line 8 and recycle gas is sent via the line 14a to the hydrodesulfurization unit 9 of step a). The hydrodesulfurization unit 9 of step a) is, for example, a reactor containing a supported hydrodesulfurization catalyst based on a group VIII metal and VIB in a fixed bed or in a fluidized bed; preferably, a fixed bed reactor is used. The reactor is operated under operating conditions and in the presence of a hydrodesulfurization catalyst as described above to decompose the sulfur compounds and to form hydrogen sulfide (H2S). During the hydrodesulfurization in step a), recombinant mercaptans are formed by addition of H2S formed to the olefins. The effluent from the hydrodesulfurization unit 9 is subsequently introduced into the “finishing” hydrodesulfurization unit 11 via the line 10 without removal of the H2S formed. Optionally, the recycle gas obtained after the step of separation via the line 14b is sent into the finishing hydrodesulfurization unit 11. The hydrodesulfurization unit 11 of step b) is, for example, a reactor containing a hydrodesulfurization catalyst in a fixed bed or in a fluidized bed; preferably, a fixed bed reactor is used. The unit 11 is operated at a higher temperature than the unit 9 and in the presence of a selective catalyst comprising an oxide support and an active phase consisting of at least one group VIII metal to decompose, at least in part, the recombinant mercaptans into olefins and into H2S. It also makes it possible to hydrodesulfurize the more refractory sulfur compounds.
According to one or more embodiments, recycle gas is sent via the line 14b into the “finishing” hydrodesulfurization unit 11, and recycle gas is optionally sent via the line 14a into the hydrodesulfurization unit 9.
An effluent (gasoline) containing H2S is withdrawn from said hydrodesulfurization reactor 11 via the line 12. The effluent is then subjected to an H2S removal step (step c) which consists, in the embodiment of
With the sequences proposed for the process according to the invention, it is possible to achieve high degrees of hydrodesulfurization while limiting the loss of olefins and consequently the decrease in the octane number. The process according to the invention also makes it possible to reduce losses of C5+ hydrocarbon compounds.
The examples that follow illustrate the invention without limiting its scope.
The analytical methods used to characterize the feedstocks and effluents are as follows:
This example makes reference to
Table 1 gives the characteristics of an FCC gasoline treated by the process according to
The FCC gasoline (line 1) is treated in the selective hydrogenation reactor 2 in the presence of a catalyst A. The catalyst A is an NiMo-on-alumina catalyst. The metal contents are respectively 7% by weight NiO and 11% by weight MoO3 with respect to the total weight of the catalyst, that is to say an Ni/Mo molar ratio of 1.2. The specific surface area of the catalyst is 230 m2/g. Prior to use thereof, the catalyst A is sulfided at atmospheric pressure in a sulfidation bed under an H2S/H2 mixture consisting of 15% by volume of H2S at 1 l/g·h of catalyst and at 400° C. for two hours. This protocol makes it possible to obtain a degree of sulfidation of greater than 80%.
The gasoline (line 1) is brought into contact with hydrogen (line 3) in a reactor which contains the catalyst A. This step of the process implements the selective hydrogenation of the diolefins and the conversion (increase in the molecular weight) of a portion of the light mercaptan (RSH) compounds present in the feedstock. The diolefin content is directly proportional to the MAV (maleic anhydride value). The diolefins are undesirable compounds since they are precursors to gums in gasolines.
The operating conditions employed in the selective hydrogenation reactor are: Temperature: 140° C., total pressure: 2.5 MPa, added H2/gasoline feedstock volume ratio: 5 normal liters of hydrogen per liter of gasoline at standard conditions (vol/vol), space velocity of the liquid: 3 h−1.
The effluent from the selective hydrogenation step (line 4) having a low content of conjugated diolefins (MAV=0.6 mg/g) and a low content of light sulfur compounds (the molecular weight of which was increased in the selective hydrogenation step) is sent to a fractionating column 5 in order to separate at the top a light gasoline (line 6) and at the bottom of the column a first heavy gasoline cut (line 7). The characteristics of the light gasoline and of the first heavy gasoline cut are indicated in table 2. As indicated in table 2, the light gasoline obtained (line 6) has a low sulfur content (10 ppm by weight). The first heavy gasoline cut, which corresponds to approximately 72% by mass of the gasoline, has a high sulfur content (600 ppm) and requires additional treatment before being incorporated into the gasoline pool.
The first heavy gasoline cut (line 7) is mixed with hydrogen (line 8) and treated in a selective hydrodesulfurization unit 9, corresponding to a first hydrodesulfurization step. The first hydrodesulfurization step is carried out in the presence of an alumina-supported CoMo catalyst, the metal contents being respectively 3% by weight CoO and 10% by weight MoO3, the specific surface area of the catalyst being 135 m2/g. Prior to use thereof, the catalyst is sulfided at atmospheric pressure in a sulfidation bed under an H2S/H2 mixture consisting of 15% by volume of H2S at 1 l/g·h of catalyst and at 400° C. for two hours. This protocol makes it possible to obtain a degree of sulfidation of greater than 80%. The temperature is 270° C., the pressure is 2.1 MPa, the liquid space velocity (expressed in volume of liquid per volume of catalyst and per hour) is 3 h−1, the ratio of the hydrogen flow rate to the feedstock flow rate is 250 normal m3 per m3 under standard conditions. The effluent from the reactor (line 10) is then reheated in an oven (not shown in the figure) and then introduced into a second reactor (11) containing a “finishing” catalyst. This finishing step is conducted in the presence of an Ni catalyst supported on alumina. The temperature is 324° C., the pressure is 1.8 MPa, the liquid space velocity (expressed in volume of liquid per volume of catalyst and per hour) is 3 h−1.
The effluent from reactor 11 (line 12) is sent to a stabilization column or debutanizer 16 operating at a pressure of 1.6 MPa, in order to separate, at the top of the column via the line 17, a stream containing H2S, hydrogen, C4-hydrocarbons and, at the bottom of the column via the line 18, a “stabilized” heavy gasoline of which the characteristics are illustrated in table 3. The loss of gasoline yield is presented in table 4.
The process as described in example 1 makes it possible to obtain a heavy gasoline having a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 35.7% by mass (in relative terms).
This example makes reference to the present invention, according to
Unless otherwise indicated below, the same feedstock as in example 1 is sent under the same operating conditions as in example 1 to the selective hydrogenation, hydrodesulfurization and finishing hydrodesulfurization sections.
The effluent from reactor 11 (line 12) is sent to a separation drum 13 after condensation at a temperature of 65° C. and a pressure of 1.6 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline and also a fraction of the dissolved H2S (line 15) and a gaseous fraction (line 14) containing essentially the hydrogen and H2S formed during steps a) and b), optionally with C1 to C4 light hydrocarbons. The liquid fraction (line 15) is then sent to a stabilization column 16 operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds from the liquid fraction and also the H2S (line 17) and a bottom cut free of H2S, called stabilized gasoline (line 18).
The gas stream (line 14) is recycled to the first hydrodesulfurization step (step a), via the line 14a.
The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in table 5 below.
The characteristics of the gasoline cut obtained after stabilization (line 18) of the present invention are illustrated in tables 6 and 7.
The process according to example 2 makes it possible to obtain a heavy gasoline having a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 34.7% by mass (in relative terms).
Unless otherwise indicated below, the same feedstock as in example 1 is sent under the same operating conditions as in example 1 to the selective hydrogenation, hydrodesulfurization and finishing hydrodesulfurization sections.
The volume flow rate of the recycle gas loop is kept constant compared to example 2; the ratio between the hydrogen flow rate and the feedstock flow rate is therefore a consequence thereof and is 125 normal m3 per m3 of the feedstock under the standard conditions, the mean temperature of the hydrodesulfurization reactor defined as being the mean between the inlet and outlet temperatures of the reactor is 283° C., the pressure is 2.2 MPa, the space velocity of the liquid (expressed as volume of liquid per volume of catalyst and per hour) is 3 h−1.
The effluent from the reactor (line 10) is reheated in an oven (not shown in the figure) and then introduced into a second reactor 11 containing a “finishing” catalyst. This finishing step is carried out at a mean temperature of 338° C., a pressure of 1.9 MPa, and a liquid space velocity (expressed in volume of liquid per volume of catalyst and per hour) is 3 h−1.
The effluent from reactor 11 (line 12) is sent to a separation drum 13 after condensation at a temperature of 65° C. and a pressure of 0.8 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline and also a fraction of the dissolved H2S (line 15) and a gaseous fraction (line 14) containing essentially the hydrogen and H2S formed during step b), optionally with C1 to C4 light hydrocarbons and C5+ hydrocarbons. The gaseous fraction is sent to an H2S separation section in an amine scrubber column 19 to obtain a recycle gas containing 10 mol ppm H2S. The recycle gas is returned, via the line 14a, as a mixture with fresh hydrogen (line 8), to the selective hydrodesulfurization unit 9 after compression.
The liquid fraction (line 15) from the separator operating at 0.8 MPa is then sent to a stabilization column 16 operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds from the liquid fraction and also the H2S (line 17) and a bottom cut free of H2S, called stabilized gasoline (line 18).
The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in table 8 below.
The characteristics of the gasoline cut obtained after stabilization (line 18) are illustrated in table 9.
The loss of C5+ compounds in the gaseous fraction obtained after the separation step (line 14) and in the gaseous cut after stabilization of the gasoline (line 17) is given in table 10.
The process according to example 3 makes it possible to obtain a heavy gasoline having a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 37.4% by mass (in relative terms). This loss of octane is greater than in the example according to the invention. This also increases hydrogen consumption in the hydrodesulfurization reactors.
Unless otherwise indicated below, the same feedstock as in example 1 is sent under the same operating conditions as in example 1 to the selective hydrogenation, hydrodesulfurization and finishing hydrodesulfurization sections.
The volume flow rate of the recycle gas loop is kept constant compared to example 2; the ratio between the hydrogen flow rate and the feedstock flow rate is therefore a consequence thereof and is 330 normal m3 per m3 of the feedstock under the standard conditions, the mean temperature of the hydrodesulfurization reactor defined as being the mean between the inlet and outlet temperatures of the reactor is 268° C., the pressure is 2.7 MPa, the space velocity of the liquid (expressed as volume of liquid per volume of catalyst and per hour) is 3 h 1.
The effluent from the reactor (line 10) is reheated in an oven (not shown in the figure) and then introduced into a second reactor (11) containing a “finishing” catalyst. This finishing step is carried out at a mean temperature of 320° C., a pressure of 2.4 MPa, and a liquid space velocity (expressed in volume of liquid per volume of catalyst and per hour) is 3 h−1.
The effluent from reactor 11 (line 12) is sent to a separation drum 13 after condensation at a temperature of 65° C. and a pressure of 2.1 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline and also a fraction of the dissolved H2S (line 15) and a gaseous fraction (line 14) containing essentially the hydrogen and H2S formed during step b), optionally with C1 to C4 light hydrocarbons and C5+ hydrocarbons. The gaseous fraction is sent to an H2S separation section in an amine scrubber column (19) to obtain a recycle gas containing 10 mol ppm H2S. The recycle gas is returned, via the line 14a, as a mixture with fresh hydrogen (line 8), to the selective hydrodesulfurization unit (9) after compression.
The liquid fraction (line 15) from the separator operating at 2.1 MPa is then sent to a stabilization column (16) operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds from the liquid fraction and also the H2S (line 17) and a bottom cut free of H2S, called stabilized gasoline (line 18).
The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in table 11.
The characteristics of the gasoline cut obtained after stabilization (line 18) are illustrated in table 12.
The loss of C5+ compounds in the gaseous fraction obtained after the separation step (line 14) and in the gaseous cut after stabilization of the gasoline (line 17) is given in table 13.
The process according to example 4 makes it possible to obtain a heavy gasoline having a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 37.8% by mass (in relative terms). This loss of octane is greater than in the example according to the invention. This also increases hydrogen consumption in the hydrodesulfurization reactors.
Unless otherwise indicated below, the same feedstock as in example 1 is sent under the same operating conditions as in example 1 to the selective hydrogenation, hydrodesulfurization and finishing hydrodesulfurization sections.
The volume flow rate of the recycle gas loop is kept constant compared to example 1; the ratio between the hydrogen flow rate and the feedstock flow rate is therefore 250 normal m3 per m3 of the feedstock under the standard conditions, the mean temperature of the hydrodesulfurization reactor is 272° C., the pressure is 3.6 MPa, the space velocity of the liquid (expressed as volume of liquid per volume of catalyst and per hour) is 3 h−1.
The effluent from the reactor (line 10) is then reheated in an oven (not shown in the figure) and then introduced into a second reactor (11) containing a “finishing” catalyst. This finishing step is carried out at a mean temperature of 330° C., a pressure of 3.3 MPa, and a liquid space velocity (expressed in volume of liquid per volume of catalyst and per hour) is 3 h−1.
The effluent from reactor 11 (line 12) is sent to a separation drum 13 after condensation at a temperature of 65° C. and a pressure of 1.6 MPa, in order to produce a liquid fraction which contains the desulfurized gasoline and also a fraction of the dissolved H2S (line 15) and a gaseous fraction (line 14) containing essentially the hydrogen and H2S formed during step b), optionally with C1 to C4 light hydrocarbons and C5+ hydrocarbons. The gaseous fraction is sent to an H2S separation section in an amine scrubber column (19) to obtain a recycle gas containing 10 mol ppm H2S. The recycle gas is returned, via the line 14a, as a mixture with fresh hydrogen (line 8), to the selective hydrodesulfurization unit (9) after compression.
The liquid fraction (line 15) from the separator operating at 1.6 MPa is then sent to a stabilization column (16) operating at a pressure of 0.7 MPa, in order to separate at the top of the column a cut of light compounds from the liquid fraction and also the H2S (line 17) and a bottom cut free of H2S, called stabilized gasoline (line 18).
The characteristics of the gaseous fraction obtained after gas/liquid separation (line 14) of the reaction effluent (line 12) are illustrated in table 14.
The characteristics of the gasoline cut obtained after stabilization (line 18) are illustrated in table 15.
The loss of C5+ compounds in the gaseous fraction obtained after the separation step (line 14) and in the gaseous cut after stabilization of the gasoline (line 17) is given in table 16.
The process according to example 5 makes it possible to obtain a heavy gasoline having a low sulfur content (10 ppm by weight). The loss of olefins between the first heavy gasoline cut and the stabilized heavy gasoline obtained after the second hydrodesulfurization step is 46.6% by mass (in relative terms). This loss of octane is greater than in the example according to the invention.
All the results are compared in table 17 below, presenting the olefin contents in gasoline and losses of C5+ compounds in the gas purges after the stabilization step. Carrying out the process according to the invention is the best compromise for maximizing the olefin content in the gasoline obtained, and for limiting the loss of C5+ compounds which can be upgraded in a gasoline cut.
| Number | Date | Country | Kind |
|---|---|---|---|
| 2114038 | Dec 2021 | FR | national |
| Filing Document | Filing Date | Country | Kind |
|---|---|---|---|
| PCT/EP2022/085356 | 12/12/2022 | WO |
