Patent Application
20120048775
In the Fischer-Tropsch process, synthesis gas (CO+H2) is catalytically transformed into oxygen-containing products and into essentially straight-chain hydrocarbons in the gas, liquid or solid form, these products constituting the feed for the process of the invention.
The paraffinic feed produced by Fischer-Tropsch synthesis used in the process of the invention is produced from a synthesis gas in the Fischer-Tropsch process; the synthesis gas (CO+H2) is advantageously produced employing three routes.
In one preferred implementation, synthesis gas (CO+H2) is produced from natural gas using the gas-to-liquid, GTL, route.
In another preferred implementation, synthesis gas (CO+H2) is produced from coal using the coal-to-liquid process, CTL.
In another preferred implementation, synthesis gas (CO+H2) is produced from biomass using the biomass-to-liquid process, BTL.
However, such products, principally constituted by normal paraffins, cannot be used as they are, in particular because their cold properties are not very compatible with the normal uses of oil cuts. As an example, the pour point of a straight-chain hydrocarbon containing 20 carbon atoms per molecule (boiling point equal to approximately 340° C., i.e. usually included in the middle distillates cut) is approximately 37° C., which renders its use impossible, since the specification for gas oil is −15° C. Thus, hydrocarbons derived from the Fischer-Tropsch process comprising mainly n-paraffins have to be transformed into more upgradeable products such as gas oil or kerosene, for example, which are obtained, for example, after catalytic hydrocracking/hydroisomerization reactions. In contrast, they may have a non-negligible quantity of unsaturated compounds of the olefinic type and oxygen-containing products (such as alcohols, carboxylic acids, ketones, aldehydes and esters). Moreover, such oxygen-containing and unsaturated compounds are concentrated in the light fractions. Thus, in the C5+ fraction corresponding to products boiling at an initial boiling point in the range 15° C. to 40° C., these compounds represent in the range 10-20% by weight of unsaturated olefinic type compounds and in the range 5-10% by weight of oxygen-containing compounds.
Such products are generally free of heteroatomic impurities such as sulphur or nitrogen, but may contain small quantities of Fe, Co, Zn, Ni or Mo originating from the dissolution of catalyst fines by the carboxylic acids. Those metals may form complexes with the oxygen-containing compounds. Said products contain no or practically no aromatics, naphthenes and more generally cycles, in particular in the case of cobalt catalysts.
The hydrogenation of unsaturated olefinic type compounds present in hydrocarbons from the Fischer-Tropsch process is a highly exothermic reaction. Thus, under the severe hydrocracking/hydroisomerization operating conditions, the transformation of said unsaturated compounds may have a negative impact on the hydrocracking step, such as thermal runaway of the reaction, substantial coking of the catalyst or the formation of gum by oligomerization. In order to protect the hydrocracking step, a hydrotreatment step is carried out under conditions which are less severe than those of hydrocracking step. However, the impurities in the feed, the oxygen-containing compounds and the metals (Fe, Co, Zn, Ni, Mo) have a deleterious effect not only on the activity of the hydrotreatment and hydrocracking catalysts, but also on the stability of the hydrotreatment catalyst. In fact, in the hydrotreatment reactor, the operating conditions as regards temperature are such that the oxygen-containing compounds do not decompose but are adsorbed onto the catalyst and form coke. In the hydrocracking section, the severe operating conditions cause the decomposition of oxygen-containing compounds into water, carbon monoxide (CO) and carbon dioxide (CO2) which are inhibitors of the acid functions (water) and the hydrogenating function (CO, CO2) of the hydrocracking catalyst and which thus modifies the activity and selectivity for middle distillates. As a consequence, the presence of alcohol or acid type oxygen-containing compounds present in the feeds necessitates an increase in the temperature of the hydrotreatment and hydrocracking step in order to compensate for the drop in activity and maintain the conversion. Further, the carboxylic acids can extract active particles from the hydrotreatment and hydrocracking catalysts, thus reducing the service life of said catalysts. Similarly, the metals complexed by the oxygen-containing compounds decompose on the active site of said hydrotreatment and hydroisomerization/hydrocracking catalysts in the presence of hydrogen and very selectively poison the active sites of said catalysts.
One of the aims of the invention is thus to reduce the total oxygen content of the feed and thus to limit the inhibiting effects of the oxygen-containing compounds and thereby limit the increase in the temperature in order to compensate for the drop in activity and maintain the conversion on the two steps, hydrotreatment and hydrocracking.
Thus, upstream of the hydrotreatment step and in order to increase the service life of the hydrotreatment catalyst and of the hydrocracking catalyst, the Applicant has instigated a step allowing transformation on an ion exchange resin, simultaneously or otherwise, of the alcohols and carboxylic acids constituting the oxygen-containing compounds into esters, and of capturing the metals complexed by said oxygen-containing compounds.
This step is followed by separation of water before the hydrotreatment step, which can reduce the total oxygen content and thus limit the inhibiting effects of the oxygen-containing compounds, and thereby limit the increase in temperature in order to compensate for the drop in activity and maintain the conversion over the two steps, hydrotreatment and hydrocracking. The water separation can also wash and capture CO and CO2, which are inhibitors, dissolved in the feed.
Shell's patent application (EP-0 583 836) describes a process for producing middle distillates from a feed obtained by the Fischer-Tropsch process. In this process, the feed from the Fischer-Tropsch synthesis may be treated in its entirety, but preferably the C4− fraction is removed from the feed so that only the C5+ fraction boiling at a temperature of over 15° C. is introduced into the subsequent step. Said feed undergoes hydrotreatment in order to hydrogenate the olefins and alcohols in the presence of a large excess of hydrogen, so that the conversion of products boiling above 370° C. into products with a lower boiling point is less than 20%. The hydrotreated effluent constituted by paraffinic hydrocarbons with a high molecular weight is preferably separated from hydrocarbon compounds with a low molecular weight, in particular the C4− fraction, before the second hydroconversion step. At least a portion of the remaining C5+ fraction then undergoes a hydrocracking/hydroisomerization step with at least 40% by weight conversion of products boiling above 370° C. into products with a lower boiling point.
Neither the presence of impurities in the feed nor the presence of steps for eliminating such impurities is mentioned in that application. Thus, Shell's patent application (EP-0 583 836) does not deal with the problem of eliminating the impurities present in the feed from the Fischer-Tropsch process.
SASOL's patent applications (WO-06/005084) and WO-06/005085 concern the elimination of metals complexed by the oxygen-containing compounds present in a paraffinic feed derived from the Fischer-Tropsch process. Those patents describe decomposition after adding water in a zone for hydrothermal conversion of those compounds. That decomposition is followed by a physical treatment which can remove the metals after decomposition. Those applications require adding water to the system, the existence of three steps (reaction, water separation, filtration) and do not affect the oxygen-containing compounds present in the feed. The present invention can reduce the number of steps required, dispense with the addition of water and simultaneously carry out transformation of the oxygen-containing compounds present in the feed.
More precisely, the present invention concerns a process for producing middle distillates from a paraffinic feed produced by Fischer-Tropsch synthesis and divided into two fractions, a light fraction, termed the cold condensate, and a heavy fraction, termed the waxes, comprising the following steps:
Throughout the remainder of the description, we shall detail the various steps in the process of the invention by referring to
At the outlet from the Fischer-Tropsch synthesis unit, the effluent derived from the Fischer-Tropsch synthesis unit is divided into two fractions, a light fraction termed the cold condensate 1, and a heavy fraction, termed the waxes, 11.
The two fractions defined thereby comprise water, carbon dioxide (CO2), carbon monoxide (CO) and unreacted hydrogen (H2). Further, the light fraction, the cold condensate, contains light C1 to C4 hydrocarbons, termed the C4− fraction, in the gaseous form.
In accordance with step a) of the process of the invention, the light fraction, termed the cold condensate, is fractionated into two fractions, a gaseous C4− fraction boiling at a temperature below 15° C. and an intermediate fraction with an initial boiling point in the range 15° C. to 40° C. and with an end point in the range 300° C. to 450° C.
The light fraction, termed the cold condensate 1, enters a fractionation means 2. The fractionation means 2 may, for example, be constituted by processes which are well known to the skilled person such as a flash drum, a distillation or a stripper. Advantageously, the fractionation means is a distillation column which can eliminate light and gaseous C1 to C4 hydrocarbons, termed the gaseous fraction C4−, 3, corresponding to products boiling at a temperature of less than 15° C., preferably less than 10° C. and more preferably less than 0° C., and recovery of an intermediate fraction 5, corresponding to C5-C30 hydrocarbons, with an initial boiling point in the range 15° C. to 40° C., preferably in the range 15° C. to 25° C., and an end point in the range 300° C. to 450° C., preferably in the range 380° C. to 400° C.
Said intermediate fraction 5 and the heavy fraction, termed the waxes, are then treated separately before being recombined in order to obtain a purified C5+ liquid fraction in line 15, corresponding to products boiling at an initial boiling point in the range 15° C. to 40° C., preferably with a boiling point of 20° C. or more.
Preferably, prior, to step b) of the process of the invention, the intermediate fraction 5 with an initial boiling point in the range 15° C. to 40° C. and an end point in the range 300° C. to 450° C. undergoes an optional decontamination step in a reactor 34, shown in
The implementations of said optional decontamination step as well as the guard bed catalysts used are the same as those used in decontamination step d) of the process of the invention and are described below in step d).
In accordance with step b) of the process of the invention, said intermediate fraction 5 with an initial boiling point in the range 15° C. to 40° C. and an end point in the range 300° C. to 450° C., optionally previously decontaminated by passage over a guard bed, passes over at least one ion exchange resin which can esterify the alcohols and carboxylic acids into esters and/or capture metals dissolved in the feed, at a temperature in the range 50° C. to 150° C., preferably in the range 80° C. to 150° C., at a total pressure in the range 0.7 to 2.5 MPa, and at an hourly space velocity in the range 0.2 to 2.5 h−1.
Step b) of the invention may advantageously be carried out in accordance with two distinct implementations, namely either in a single reactor 6 over a single ion exchange resin, advantageously used to simultaneously carry out esterification of alcohols and carboxylic acids to esters and capture of the metals dissolved in the feed, or in two different reactors (not shown in the figures) over two ion exchange resins of different natures, one having the specific function of esterification of alcohols and carboxylic acids and the other the capture of the metals dissolved in the feed.
In a first implementation, step b) advantageously consists of passing said intermediate fraction 5 over a single ion exchange resin in a single reactor 6 to simultaneously carry out the esterification of alcohols and carboxylic acids to esters and the capture of metals dissolved in the feed.
Preferably, said resin is used at a temperature in the range 80° C. to 150° C. and preferably in the range 80° C. to 130° C., at a pressure in the range 1 to 2 MPa and preferably in the range 1 to 1.5 MPa, and at an hourly space velocity in the range 0.5 to 2 h−1, preferably in the range 0.5 to 1.5 h−1.
In this case, oxygen-containing compounds, carboxylic acids and alcohols are adsorbed onto the active sites of said resin and are esterified and the cationic and metallic compounds present in said C5+ liquid paraffinic fraction are eliminated by adsorption or by ion exchange. Said resin, which can simultaneously carry out the esterification of alcohols and carboxylic acids to esters and the capture of metals dissolved in the feed, advantageously comprises sulphonic acid groups and is prepared by polymerization or co-polymerization of aromatic vinyl groups followed by sulphonation, said aromatic vinyl groups being selected from styrene, vinyl toluene, vinyl naphthalene, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinyl xylene, said resin having a degree of cross-linking in the range 20% to 35%, preferably in the range 25% to 35%, preferably equal to 30%, and an acid strength, assayed by potentiometry during neutralization with a KOH solution, of 0.2 to 6 mmol H+ equivalent/g, preferably in the range 0.2 to 2.5 mmol H+ equivalent/g.
Said acid ion exchange resin advantageously contains in the range 1 to 2 terminal sulphonic groups per aromatic group. Preferably, said resin has a size in the range 0.15 to 1.5 mm. The size of a resin is the diameter of the sphere encompassing the resin particle. The resin size categories are measured by screening through suitable screens, in accordance with a technique which is known in the art.
A preferred resin is a resin constituted by co-polymers of monovinyl aromatics and polyvinyl aromatics, and highly preferably, a copolymer of divinyl benzene and polystyrene with a degree of cross-linking in the range 20% to 35%, preferably in the range 25% to 35%, and more preferably equal to 30%, and an acid strength, representing the number of active sites of said resin, assayed by potentiometry during neutralization using a KOH solution, in the range 0.2 to 6 mmol H+ equivalent/g, preferably in the range 0.2 to 2.5 mmol H+ equivalent/g.
Another preferred resin which can simultaneously allow the esterification of alcohols and carboxylic acids and the capture of metals to be carried out is a resin constituted by a polysiloxane grafted with alkylsulphonic type acid groups (of the —CH2—CH2—CH—2—SO3H type), with a size in the range 0.5 to 1.2 mm and with an acid strength, representing the number of active sites of said resin and assayed by potentiometry during neutralization with a KOH solution, of 0.4 to 1.5 mmol H+ equivalent/g.
During said step and under these conditions, 95% of the carboxylic acids are esterified. The acid conversion is analyzed by the potassium hydroxide titration difference between the feed and the effluent using a technique which is known to the skilled person. The ASTM methods D 664, D 3242 or D 974 can be cited, for example, as methods for carrying out said analysis.
This resin may advantageously be used in a fixed bed between screens placed in an upflow or downflow tube reactor. Preferably, said resin is used in an upflow bed reactor, the liquid being injected into the bottom of the reactor at a sufficient surface velocity to allow the bed of resin to expand without, however, either transporting or fluidizing it. This implementation, compared with a fixed bed, can attenuate the effects, of clogging materials and substantially increase the service life of the resin.
In accordance with a second implementation, step b) advantageously consists of passing said intermediate fraction 5 into two different reactors over two distinct ion exchange resins, of different natures, one having the specific function of esterification of alcohols and carboxylic acids and the other that of capturing the metals dissolved in the feed.
Preferably, the reactor containing the ion exchange resin which can capture metals is used upstream of the reactor containing the ion exchange resin which can carry out the esterification of the alcohols and carboxylic acids.
In this case, the cationic and metallic compounds present in said intermediate fraction 5 are eliminated by adsorption or by ion exchange on a first ion exchange resin. This first resin, which is specific for capturing metals, advantageously comprises sulphonic acid groups and is advantageously prepared by polymerization or co-polymerization of aromatic vinyl groups followed by sulphonation, said aromatic vinyl groups advantageously being selected from styrene, vinyl toluene, vinyl naphthalene, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinyl xylene, said resin having a degree of cross-linking in the range 1% to 20%, preferably in the range 2% to 8%, and an acid strength, representing the number of active sites in said resin, assayed by potentiometry during neutralization with a KOH solution, in the range 1 to 15 mmol H+ equivalent/g, preferably in the range 2.5 to 10 mmol equivalent/g.
Said acid ion exchange resin advantageously contains between 1 and 2 terminal sulphonic acid groups per aromatic group. Preferably, said resin has a size in the range 0.15 to 1.5 mm. The size of the resin is the diameter of the sphere encompassing the resin particle, The size classes for the resin are measured by screening through suitable screens using a technique which is known to the skilled person.
Preferably, the first resin is a resin constituted by copolymers of monovinyl aromatics and polyvinyl aromatics; more preferably, a copolymer of divinyl benzene and polystyrene with a degree of cross-linking in the range 1% to 20% and an acid strength, representing the number of active sites of said resin and assayed by potentiometry during neutralization with a KOH solution, of 1 to 15 mmol H+ equivalent/g, preferably in the range 2.5 to 10 mmol H+ equivalent/g.
Preferably, said first resin is used at a temperature in the range 50° C. to 110° C., preferably in the range 80° C. to 110° C., at a pressure in the range 1 to 2 MPa and preferably in the range 1 to 1.5 MPa, and at an hourly space velocity in the range 0.2 to 1.5 h−1, preferably in the range 0.5 to 1.5 h−1.
The effluent from the reactor containing said first resin which is specific for the capture of metals is then introduced into a second reactor located downstream of the first reactor and containing a second resin with a different nature and which is specific to esterification of the alcohols and carboxylic acids contained in said effluent.
The oxygen-containing compounds, carboxylic acids and alcohols are adsorbed onto the active sites of said second resin and are esterified and the cationic and metallic compounds present in said intermediate fraction 5 are eliminated by adsorption or by ion exchange. Said second resin, which can simultaneously carry out the esterification of alcohols and carboxylic acids to esters advantageously comprises sulphonic acid groups and is advantageously prepared by polymerization or copolymerization of aromatic vinyl groups followed by sulphonation. The aromatic vinyl groups are advantageously selected from styrene, vinyl toluene, vinyl naphthalene, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinyl xylene, said second resin having a degree of cross-linking, i.e. a ratio of the mass of copolymer/mass of polymer, which is advantageously in the range 20% to 35%, preferably in the range 25% to 35% and more preferably 30%, and an acid strength, representing the number of active sites of said resin, assayed by potentiometry during neutralization with a KOH solution, in the range 0.2 to 6 mmol H+ equivalent/g, preferably in the range 0.2 to 6 mmol H+ equivalent/g.
Said second acid ion exchange resin advantageously contains 1 to 2 terminal sulphonic groups per aromatic group. Preferably, said second resin has a size in the range 0.15 to 1.5 mm.
A preferred second resin is a resin constituted by copolymers of monovinyl aromatics and aromatic polyvinyls, and more preferably, a copolymer of divinyl benzene and polystyrene with a degree of cross-linking in the range 20% to 35%, preferably in the range 25% to 35% and more preferably 30%, and an acid strength, representing the number of active sites of said resin, assayed by potentiometry during neutralization with a KOH solution, in the range 0.2 to 6 mmol H+ equivalent/g and preferably in the range 0.2 to 6 mmol H+ equivalent/g.
A further preferred second resin which can simultaneously allow esterification of alcohols and carboxylic acids and capture of metals to be carried out is a resin constituted by a polysiloxane grafted with alkylsulphonic type acid groups (of the —CH2—CH2—CH2—SO3H type), with a size in the range 0.5 to 1.2 mm and with an acid strength, representing the number of active sites of said resin and assayed by potentiometry during neutralization with a KOH solution, of 0.4 to 1.5 mmol H+ equivalent/g.
Preferably, said second resin is used at a temperature in the range 80° C. to 150° C. and preferably in the range 80° C. to 130° C., at a pressure in the range 1 to 2 MPa and preferably in the range 1 to 1.5 MPa, and at an hourly space velocity in the range 0.5 to 2 h−1, preferably in the range 0.5 to 1.5 h−1.
During this step and under these conditions, 95% of the carboxylic acids are esterified. Analysis of the conversion of the acids is given by the difference in the potassium hydroxide titration between the feed and the effluent. The ASTM methods D 664, D 3242 or D 974 can be cited, for example, as methods for carrying out said analysis. These resins may advantageously be used in a fixed bed between screens placed in an upflow or downflow tube reactor. Preferably, said resin is used in an upflow bed reactor, the liquid being injected into the bottom of the reactor at a sufficient surface velocity to allow the bed of resin to expand without, however, either transporting or fluidizing it. This implementation, compared with a fixed bed, can attenuate the effects of clogging materials and substantially increase the service life of the resin.
In the case in which said intermediate fraction 5 passes into two different reactors over two distinct ion exchange resins of different natures, one principally carrying out metals capture, the other principally carrying out esterification, intermediate re-heating is preferably employed between the two steps. Preferably, the water formed during the step for esterification of the acids and alcohols over the first resin principally carrying out the capture of metals is removed in order to intensify the esterification reaction over the second resin. Simultaneously adding intermediate re-heating and water separation boosts the overall conversion of the carboxylic acids present in the feed.
The reaction for esterification of the organic acids by the alcohols present in said intermediate fraction 5 produces water which is a compound that inhibits the hydrotreatment and hydrocracking catalysts, necessitating an increase in the severity of the operating conditions.
In accordance with the invention, the effluent derived from step b) then undergoes a step for eliminating at least a portion of the water formed during said step b), preferably all of the water formed, in a separator 8.
This water is acidic in nature as it advantageously contains protons exchanged during capture of the metals by the upstream specific cation exchange resin or by the only resin allowing simultaneous esterification and, capture of metals. This water may also contain dissolved CO and CO2 originating from the Fischer-Tropsch synthesis. The water is eliminated via the line 9.
It is also advantageous to add in said step c) gas of the nitrogen (N2) or hydrogen (H2) type, to eliminate more dissolved CO and CO2 by stripping.
In the case when hydrogen is added, this advantageously acts as a makeup gas for the hydrotreatment step.
This step can also eliminate products of the light ether type formed during the reaction of alcohols with themselves.
The water may be eliminated using any of the methods and techniques known to the skilled person, for example by drying, passage over a dessicant, flash, decantation, etc.
The effluent from water elimination step c) is then recombined, in accordance with step e) of the process of the invention, with the effluent from step d) described below.
Step d)
In accordance with the invention, the heavy fraction termed the waxes undergoes a step for decontamination in a reactor 12 by passage over a guard bed containing at least one guard bed catalyst.
The treated heavy fraction may possibly contain solid particles such as mineral solids. They may possibly contain metals contained in hydrocarbon structures such as organometallic compounds of greater or lesser solubility. The term “fines” means fines resulting from physical or chemical attrition of the catalyst. They may be on the micron or sub-micron scale. These mineral particles thus contain the active components of these catalysts; a non-limiting list is as follows: alumina, silica, titanium, zirconia, cobalt oxide, iron oxide, tungsten, ruthenium oxide, etc. These solid minerals may be in the form of a calcined mixed oxide: examples are alumina-cobalt, alumina-iron, alumina-silica, alumina-zirconia, alumina-titanium, alumina-silica-cobalt, alumina-zirconia-cobalt, etc.
Said heavy fraction may also contain metals within hydrocarbon structures, which may possibly contain oxygen or organometallic compounds of greater or lesser solubility. More particularly, said compounds may be silicon-based. As an example, they may be anti-foaming agents used in the synthesis process. As an example, the solutions of a silicone type silicon compound or silicone oil emulsion are more particularly contained in the heavy fraction.
Further, the catalyst fines described above may have a silica content which is greater than the formulation for the catalyst, resulting from intimate interaction between the catalyst fines and the anti-foaming agents described above.
The problem which thus arises is to reduce the quantity of solid mineral particles and possibly to reduce the quantity of metallic compounds which are deleterious to the hydroisomerization-hydrocracking catalyst.
The guard beds advantageously contain at least one catalyst.
The guard bed catalysts used in steps c) of the process of the invention and in the optional step for decontamination of said intermediate fraction may be different and are preferably identical. Said guard bed catalysts are described below.
The catalysts in the guard beds used in steps c) of the process of the invention and in the optional step for decontamination of said intermediate fraction may advantageously have the shape of spheres or extrudates. However, it is advantageous for the catalyst to be in the shape of extrudates with a diameter in the range 0.5 to 5 mm, more particularly in the range 0.7 to 2.5 mm. The shapes are cylinders (which may or may not be hollow), twisted cylinders, multilobes (2, 3, 4 or 5 lobes, for example), or rings. The cylindrical shape is preferred, but any other shape may be used.
To accommodate the presence of contaminants and/or poisons in the feed, in a further preferred implementation, the guard bed catalysts may have more particular geometrical forms to increase their void fraction. The void fraction of said catalysts is in the range 0.2 to 0.75. Their external diameter may be between 1 and 35 mm. Possible particular non-limiting forms are: hollow cylinders, hollow rings, Raschig rings, toothed hollow cylinders, crenellated hollow cylinders, pentaring cartwheels, multiple holed cylinders, etc.
Said catalysts of the guard beds used in steps c) of the process of the invention and in the optional step for decontamination of said intermediate fraction may advantageously have been impregnated with a phase which may or may not be active. Preferably, the catalysts are impregnated with a hydrodehydrogenating phase. Highly preferably, the CoMo or NiMo phase is used. Still more preferably, the NiMo phase is used.
Preferably, the supports for said guard bed catalysts are porous refectory oxides, preferably selected from alumina and silica-alumina.
Said guard bed catalysts may advantageously have macroporosity.
Said catalysts advantageously have a macroporous mercury volume for a mean diameter of 50 nm which is more than 0.1 cm3/g, preferably in the range 0.125 to 0.175 cm3/g, and a total volume of more than 0.60 cm3/g, preferably in the range 0.625 to 0.8 cm3/g, and is advantageously impregnated with an active phase, preferably based on nickel and molybdenum, such as ACT961, for example. In this preferred embodiment, the Ni content as the weight of oxide is generally in the range 1% to 10% and the Mo content as the weight of oxide is in the range 5% to 15%. The surface areas, expressed as the SBBT, of the supports for said catalysts are in the range 30 m2/g to 220 m2/g.
In a first embodiment, the guard bed advantageously also comprises at least one other catalyst having a mercury volume for a pore diameter of more than 1 micron of more than 0.2 cm3/g and preferably more than 0.5 cm3/g, and a mercury volume for a pore diameter of more than 10 microns of more than 0.25 cm3/g and preferably less than 0.4 cm3/g, said catalyst advantageously being placed upstream of the catalyst of the invention.
In a second embodiment, the guard bed advantageously also comprises at least one other catalyst with a mercury volume for a pore diameter of more than 50 nm of more than 0.25 cm3/g, the mercury volume for a pore diameter of more than 100 nm being more than 0.15 cm3/g and a total pore volume of more than 0.80 cm3/g.
Said guard bed catalyst and the catalyst of the first embodiment may advantageously be associated in a mixed bed or a combined bed. In general, the impregnated catalyst of the active phase constitutes the majority of the guard bed and the catalyst of the first embodiment which is preferred is added as a complement of 0 to 50% by volume with respect to the first catalyst, preferably 0 to 30%, more preferably 1% to 20%.
The combination of said guard bed catalyst and the catalyst of the first embodiment does not limit the scope of the invention. The catalysts which can be used in the guard beds may advantageously be used alone or as a mixture; in a non-exhaustive manner, they may be selected from those sold by Norton-Saint-Gobain, for example MacroTrap® guard beds, or catalysts sold by Axens from the ACT family: ACT077, ACT935, ACT961 or HMC841, HMC845, HMC941 or HMC945.
Preferred guard beds for use in the invention are the HMCs and ACT961. It may be particularly advantageous to superimpose these catalysts in at least two different beds of varying heights. The catalysts with the highest void ratio are preferably used in the first catalytic bed or beds at the inlet to the catalytic reactor. It may also be advantageous to use at least two different reactors for said catalysts.
Advantageously, an association of said guard bed catalyst with the catalysts of the first and second implementation is also possible in a mixed bed or a combined bed. In this case, the catalysts are placed with the void capacity decreasing in the direction of flow.
After passing over said guard bed, the quantity of solid particles is less than 20 ppm, preferably less than 10 ppm and more preferably less than 5 ppm. The soluble silicon content is less than 5 ppm, preferably less than 2 ppm and more preferably less than 1 ppm.
Preferably, the effluent derived from decontamination step d) by passing a heavy fraction termed the waxes over a guard bed can optionally pass, prior to the step e) for recombination of the purified intermediate fraction derived from step e) (line 10) and said effluent derived from step d), over at least one ion exchange resin at a temperature in the range 50° C. to 150° C., preferably in the range 80° C. to 150° C., at a total pressure in the range 0.7 to 2.5 MPa, at an hourly space velocity in the range 0.2 to 2.5 h−1.
This optional decontamination step may advantageously be carried out in accordance with two distinct implementations, namely either in a single reactor 35 (shown in
The above implementations as well as the resins used are the same as those used in step b) of the process of the invention and have been described above.
In accordance with the invention, the purified intermediate fraction derived from step c) (line 10) and the effluent derived from decontamination step d), optionally purified by passage over at least one ion exchange resin (line 13), are recombined in the line 15 in order to obtain a purified C5+ fraction which constitutes the feed for step f) for hydrogenation of unsaturated olefinic type compounds.
Step f) of the process of the invention is a step for hydrogenation of the unsaturated olefinic type compounds of at least a portion and preferably the whole of the effluent derived from step e) of the process of the invention, in the presence of hydrogen and a hydrogenation catalyst.
The effluent from step e) (line 15) of the process of the invention is admitted in the presence of hydrogen (line 14) into a hydrogenation zone 16 containing a hydrogenation catalyst which is intended to saturate the unsaturated olefinic type compounds present in said effluent.
Preferably, the catalyst used in step f) of the invention is a non-cracking or low cracking hydrogenation catalyst comprising at least one metal from group VIII of the periodic table of the elements and comprising at least one support based on a refractory oxide.
Preferably, said catalyst comprises at least one metal from group VIII selected from nickel, cobalt, ruthenium, indium, palladium and platinum and comprising at least one support based on refractory oxide selected from alumina and silica-alumina.
Preferably, the metal from group VIII is selected from nickel, palladium and platinum;
highly preferably, it is selected from palladium and platinum.
In accordance with a preferred implementation of step f) of the process of the invention, the metal from group VIII is selected from palladium and/or platinum and the quantity of this metal is advantageously in the range 0.1% to 5% by weight, preferably in the range 0.2% to 0.6% by weight with respect to the total catalyst weight.
In accordance with a highly preferred implementation of step f) of the process of the invention, the metal from group VIII is palladium.
According to another preferred implementation of step f) of the process of the invention, the metal from group VIII is nickel and the quantity of this metal is advantageously in the range 5% to 25% by weight, preferably in the range 7% to 20% by weight with respect to the total catalyst weight.
The support for the catalyst used in step f) of the process of the invention is a support based on a refractory oxide, preferably selected from alumina and silica-alumina, more preferably alumina.
When the support is an alumina, it has a BET specific surface area which can limit polymerization reactions at the surface of the hydrogenation catalyst, said surface area being in the range 5 to 140 m2/g.
When the support is a silica-alumina, the support contains a percentage of silica in the range 5% to 95% by weight, preferably in the range 10% to 80%, more preferably in the range 20% to 60% by weight and highly preferably in the range 30% to 50%, a BET specific surface area in the range 100 to 550 m2/g, preferably in the range 150 to 500 m2/g, more preferably less than 350 m2/g and still more preferably less than 250 m2/g.
The hydrogenation step f) of the process of the invention is preferably carried out in one or more fixed bed reactors.
In hydrogenation zone 16, the feed is brought into contact with the hydrogenation catalyst in the presence of hydrogen and at operating temperatures and pressures which allow hydrogenation of the unsaturated olefinic type compounds present in the feed.
The operating conditions for hydrogenation step f) of the process of the invention are advantageously as follows: the temperature in said hydrogenation zone 16 is in the range 100° C. to 180° C., preferably in the range 120° C. to 165° C., the total pressure is in the range 0.5 to 6 MPa, preferably in the range 1 to 5 MPa and more preferably in the range 2 to 5 MPa. The flow rate of the feed is such that the hourly space velocity (ratio of the hourly volume flow rate at 15° C. for fresh liquid feed to the volume of charged catalyst) is in the range 1 to 50 Nl/l/h, preferably in the range 2 to 20 h−1 and more preferably in the range 4 to 20 h−1. The hydrogen which supplies the hydrotreatment zone is introduced at a flow rate such that the hydrogen/hydrocarbon volume ratio is in the range 5 to 300 Nl/l/h, preferably in the range 5 to 200, more preferably in the range 10 to 150 Nl/l/h, and still more preferably in the range 10 to 50 Nl/l/h.
Under these conditions, the unsaturated olefinic type compounds are more than 50%, preferably more than 75% and more preferably more than 85% hydrogenated.
The effluent from step f) optionally undergoes a step for elimination of at least a portion of the water formed during hydrogenation step f), preferably all of the water which is formed, in a separator 37 shown in
This water may also contain a fraction of dissolved CO and CO2 originating from the Fischer-Tropsch synthesis. This step for eliminating water takes place in the separator 37 and water is eliminated via the line 39.
It may also be advantageous to add to said step for elimination of at least a portion of the water a nitrogen (N2) or hydrogen (H2) type gas to eliminate more dissolved CO and CO2 by stripping.
This step can also eliminate light ether type products formed during the reaction of alcohols on themselves.
The water may be eliminated using any of the methods and techniques known to the skilled person, for example drying, passage over a dessicant, flash, decantation, etc.
At the end of step f) of the process of the invention, at least a portion and preferably all of the liquid hydrogenated effluent is sent to a hydrocracking/hydroisomerization zone 19.
In accordance with step g) of the process of the invention, at least a portion and preferably all of the liquid hydrogenated effluent from hydrogenation step f) of the process of the invention is sent to the hydroisomerization/hydrocracking zone 19 containing the hydroisomerization/hydrocracking catalyst, preferably at the same time as a stream of hydrogen.
The operating conditions in which hydroisomerization/hydrocracking step g) of the process of the invention is carried out are preferably as follows:
The pressure is generally maintained between 0.2 and 15 MPa, preferably in the range 0.5 to 10 MPa and advantageously in the range 1 to 9 MPa; the hourly space velocity is generally in the range 0.1 h−1 to 10 h−1, preferably in the range 0.2 to 7 h−1 and advantageously in the range 0.5 to 5.0 h−1. The hydrogen ratio is generally in the range 100 to 2000 normal litres of hydrogen per litre of feed per hour, preferably in the range 150 to 1500 normal litres of hydrogen per litre of feed per hour.
The temperature used in this step is generally in the range 200° C. to 450° C., preferably in the range 250° C. to 450° C., advantageously in the range 300° C. to 450° C., and more advantageously more than 320° C. or, for example, in the range 320° C. to 420° C.
Hydroisomerization and hydrocracking step g) of the process of the invention is advantageously carried out under conditions such that the conversion per pass of products with a boiling point of 370° C. or more into products with boiling points of less than 370° C. is more than 80% by weight, and more preferably at least 85%, preferably more than 88%, in order to obtain middle distillates (gas oil and kerosene) with sufficiently good cold properties (pour point, freezing point) so that they satisfy specifications in force for this type of fuel.
The majority of the catalysts in current use in hydroisomerization/hydrocracking are bi-functional in type, associating an acid function with a hydrogenating function. The acid function is supplied by supports with large surface areas (generally of 150 to 800 m2/g) and with a superficial acidity, such as halogenated aluminas (in particular chlorinated or fluorinated), phosphorus-containing aluminas, combinations of oxides of boron and aluminium, and silica-aluminas. The hydrogenating function is generally supplied either by one or more metals from group VIII of the periodic table of the elements such as iron, cobalt; nickel, ruthenium, rhodium, palladium, osmium; iridium or platinum, or by a combination of at least one metal from group VI, such as chromium, molybdenum or tungsten, and at least one group VIII metal.
In the case of bi-functional catalysts, the balance between the two functions, acid and hydrogenating, is the fundamental parameter which governs the activity and selectivity of the catalyst. A weak acid function and a strong hydrogenating function produces less active catalysts which are also less selective as regards isomerization, while a strong acid function and a weak hydrogenating function produces catalysts which are highly active and selective as regards cracking. A third possibility is to use a strong acid function and a strong hydrogenating function to obtain a catalyst which is highly active but also highly selective as regards isomerization. Thus, by carefully selecting each of the functions, it is possible to adjust the activity/selectivity balance of the catalyst.
Advantageously, the hydroisomerization/hydrocracking catalysts are bi-functional catalysts comprising an amorphous acid support (preferably a silica-alumina) and a metallic hydro-dehydrogenating function which is preferably provided by at least one noble metal. The support is termed amorphous, i.e. free of molecular sieve and in particular zeolite, as is the catalyst. The amorphous acid support is advantageously a silica-alumina, but other supports may be used. When it is a silica-alumina, the catalyst preferably contains no added halogen other than that which may be introduced for impregnation with the noble metal, for example.
More generally and preferably, the catalyst contains no added halogen, for example fluorine. In general and preferably, the support has not undergone impregnation with a silicon compound.
In accordance with a first preferred implementation, the hydroisomerization/hydrocracking catalyst contains at least one hydrodehydrogenating element selected from noble group VIII metals, preferably platinum and/or palladium, and at least one amorphous refractory oxide support, preferably silica-alumina.
A preferred hydroisomerization/hydrocracking catalyst used in step g) of the process of the invention comprises up to 3% by weight of a metal of at least one hydro-dehydrogenating element selected from noble metals from group VIII, preferably deposited on the support, and highly preferably, the noble group VIII metal is platinum; and a support comprising (or preferably constituted by) at least one silica-alumina, said silica-alumina having the following characteristics:
A second preferred hydroisomerization/hydrocracking catalyst used in step g) of the process of the invention comprises up to 3% by weight of a metal of at least one hydro-dehydrogenating element selected from noble metals from group VIII of the periodic table of the elements, and preferably, the noble group VIII metal is platinum; 0.01% to 5.5% by weight of oxide of a doping element selected from phosphorus, boron and silicon; and a non-zeolitic support based on silica-alumina containing a quantity of more than 15% by weight and 95% by weight or less of silica (SiO2), said silica-alumina having the following characteristics:
Advantageously, the characteristics associated with the corresponding catalyst are identical to those of the silica-alumina described above.
The two steps f) and g) of the process of the invention, hydrogenation and hydroisomerization-hydrocracking, may advantageously be carried out on the two types of catalysts in two or more different reactors and/or in the same reactor.
In accordance with a second preferred implementation, the hydroisomerization/hydrocracking catalyst contains at least one hydrodehydrogenating element selected from non-noble group VIII metals and metals from group VIB and at least one amorphous refractory oxide support, preferably silica-alumina.
Preferably, the metal from group VIII is selected from nickel and cobalt, and the metal from group VIB is selected from molybdenum and tungsten.
Preferably, said catalyst is in the sulphide form.
A third preferred hydroisomerization/hydrocracking catalyst used in step g) of the process of the invention comprises at least one hydro-dehydrogenating element selected from non-noble metals from group. VIII and metals from group VIB of the periodic table of the elements, preferably between 2.5% and 5% by weight of oxide of the non-noble element from group VIII and between 20% and 35% by weight of oxide of a group VIB element with respect to the weight of the final catalyst; preferably, the non-noble group VIII metal is nickel and the group VIB metal is tungsten; optionally 0.01% to 5.5% by weight of oxide of a doping element selected from phosphorus, boron and silicon; preferably, 0.01% to 2.5% by weight of oxide of a doping element and a non-zeolitic support based on silica-alumina containing a quantity of more than 15% by weight and 95% by weight or less of silica (SiO2), preferably a quantity of more than 15% by weight and 50% by weight or less of silica, said silica-alumina having the following characteristics:
Advantageously, the characteristics associated with the corresponding catalyst are identical to those of the silica-alumina described above.
When the third preferred hydroisomerization/hydrocracking catalyst is used in step g) of the process of the invention, said catalyst is sulphurized.
In accordance with a first preferred implementation of the process of the invention, in hydrogenation step e), a catalyst is used which contains palladium and in hydroisomerization/hydrocracking step g), a catalyst containing platinum is used.
In accordance with a second preferred implementation of the process of the invention, in hydrogenation step f) a catalyst containing palladium is used and in hydroisomerization/hydrocracking step g), a sulphurized catalyst containing at least one hydro-dehydrogenating element selected from non-noble metals from group VIII and group VIB metals is used.
In a third preferred implementation of the process of the invention, in hydrogenation step f) a catalyst containing at least one non-noble hydro-dehydrogenating element from group VIII is used and in hydroisomerization/hydrocracking step g), a sulphurized catalyst containing at least one hydro-dehydrogenating element selected from non-noble group VIII metals and group VIB metals is used.
In accordance with step h) of the process of the invention, the effluent derived from step g) (line 20) undergoes the separation of unreacted hydrogen and light gases in a gas/liquid separator 33 then recycling of the unreacted hydrogen and said light gases to the hydroisomerization/hydrocracking step g) (line 22).
Said light gases include light C1-C4 gases, carbon monoxide (CO), carbon dioxide (CO2) and water in the vapour form.
Said gases are separated from the liquid effluent in one or more flash drums, i.e. one or more drums which carry out separation of the gases and the liquids introduced via the line 20, at staged temperatures and pressures in order to increase the recovery of hydrogen. This flash staging may advantageously be accompanied by a heat exchanger aimed at recovering heat energy and/or cooling the effluents from the separator drums in order to minimize losses of hydrogen.
By using an ion exchange resin upstream of the hydrotreatment and hydrocracking steps, the process of the invention can reduce the total oxygen content of the feed and thus limit the formation of carbon monoxide (CO) originating from the decomposition of oxygen-containing compounds present in the feed in the hydroisomerization/hydrocracking section. Carbon monoxide (CO) is an inhibitor of the metallic compounds present on the hydroisomerization/hydrocracking catalyst, and its content must be minimized in order not to require an increase in temperature in order to compensate for the low activity and maintain conversion.
However, when the gaseous effluent 22 derived from said separation has a high CO fraction, i.e. more than 10 ppm by volume, said gaseous effluent is advantageously sent to a methanation reactor 40, shown in
A purge is advantageously carried out (line 23) in order to eliminate the products formed during the methanation step 33. A makeup of hydrogen (line 24) is then advantageously carried out in order to compensate for that purge.
In accordance with step i) of the process of the invention, the effluent from step h) for separating unreacted hydrogen and the light gases of the process of the invention is sent, to a distillation train 26 via a line 21, which combines atmospheric distillation with optional vacuum distillation, which is intended to separate conversion products with a boiling point of less than 340° C. and preferably less than 370° C. and in particular including those formed during step g) in the hydroisomerization/hydrocracking reactor 19, and to separate the residual fraction with an initial boiling point which is generally more than at least 340° C. and preferably at least 370° C. or higher. Of the converted and hydroisomerized products, in addition to the light C1-C4 gases (line 27), at least one gasoline (or naphtha) fraction is separated (line 28), and at least one kerosene middle distillate fraction (line 29) and a gas oil fraction (line 30) are separated. Preferably, the residual fraction, with an initial boiling point which is generally over at least 340° C. and preferably at least 370° C. is recycled (line 17) to step g) of the process of the invention to the head of the hydroisomerization and hydrocracking zone 19. In accordance with another implementation of step i) of the process of the invention, said residual fraction may supply excellent oil bases.
It may also be advantageous to recycle (line 32) at least part and preferably all of at least one of the kerosene and gas oil cuts obtained to step g) (line 19). The gas oil and kerosene cuts are preferably recovered separately or mixed, but the cut points are adjusted by the operator as a function of requirements. It has been shown that it is advantageous to recycle part of the kerosene to improve its cold properties.
The gas oil(s) obtained have a pour point of at most 0° C., generally less than −10° C. and usually less than −15° C. The cetane index is more than 60, generally more than 65, and usually more than 70.
The kerosene(s) obtained have a freezing point of at most −35° C., generally less than −40° C. The smoke point is more than 25 mm, generally more than 30 mm. In this process, gasoline (unwanted) production is as low as possible. The gasoline yield is always less than 50% by weight, preferably less than 40% by weight, advantageously less than 30% by weight or 20% by weight or even 15% by weight.
At the outlet from the Fischer-Tropsch unit, the effluent from the Fischer-Tropsch synthesis unit was divided into two fractions, a light fraction termed the cold condensate and a heavy fraction, termed the waxes.
The light fraction, termed the cold condensate, was fractionated into a gaseous C4− fraction boiling at a temperature of less than 15° C. and an intermediate fraction with an initial boiling point in the range 15° C. to 40° C. and an end point in the range 300° C. to 450° C. The characteristics of the various fractions are given in Table 1 below:
The intermediate C5-C30 fraction passed over an ion exchange resin with trade name Amberlyst 35 sold by Röhm & Haas, said resin allowing simultaneous capture of metals dissolved in the feed and esterification of alcohols and carboxylic acids to esters. Said resin was constituted by divinyl benzene-polystyrene copolymers with a degree of cross-linking of 20 and an acid strength, assayed by potentiornetry during neutralization with a KOH solution, of 4.15 mmol H+ equivalent/g.
Step b) was carried out at a temperature of 100° C., a pressure of 1 MPa, and at an hourly space velocity of 1 h−1. Under these conditions, 95% of the acids were esterified, the analysis of the conversion of the acids being given by the difference in potassium hydroxide titration between the feed and the effluent using the ASTM D664 method. The composition of the outlet effluent is given in Table 2.
The effluent derived from step c) then underwent the elimination of the water formed during said step b), by decanting/coalescence in suitable equipment known to the skilled person. The effluent at the outlet contained 300 ppm of water.
The heavy fraction, termed the waxes, derived from the Fischer-Tropsch synthesis unit (line 11) passed over a guard bed composed of ACT 961 sold by Axens, at a temperature of 80° C., a pressure of 1 MPa, an hourly space velocity of 1 h−1. At the outlet, the effluent (line 13) contained less than 1 ppm of fines.
The purified intermediate fraction (line 10) and the wax fraction derived from step d) (line 13) were mixed in their entirety. The composition of the mixture formed thereby is given in Table 3 below:
All of the effluent derived from step e) then underwent a step for hydrogenation in the presence of hydrogen and a hydrogenation catalyst with trade name LD265 sold by Axens, said catalyst comprising 0.3% by weight of palladium deposited on an alumina with a specific surface area of 69 m2/g.
Hydrogenation step f) was carried out at a reaction temperature of 130° C., at a pressure of 3.5 MPa, the hydrogen was introduced at a flow rate such that the hydrogen/hydrocarbon volume ratio was 32 Nl/l/h, and with an hourly space velocity of 10 n−1. Under these conditions, the conversion of olefins was 85% by weight.
The liquid effluent derived from hydrogenation step f) had the composition described in Table 4:
All of the effluent from hydrogenation step f) underwent a hydroisomerization/hydrocracking step in the presence of fresh hydrogen and a hydroisomerization/hydrocracking catalyst, in which the residual fraction with an initial boiling point of more than 370° C., unreacted hydrogen and light gases were recycled.
The hydroisomerization/hydrocracking catalyst comprised 0.6% by weight of platinum and a support comprising 29.3% by weight of silica, SiO2, and 70.7% by weight of alumina, Al2O3, a Na content of 100 ppm by weight, a total pore volume comprising 0.69 ml/g measured by mercury porosimetry, a volume of mesopores with a mean diameter of 80 Å representing 78% of the total pore volume measured by mercury porosimetry, a volume of macropores with a diameter of more than 500 Å representing 22% of the total pore volume measured by mercury porosimetry and a specific surface area of 300 m2/g.
The hydroisomerization/hydrocracking step was carried out under the conditions described in Table 5.
The conversion per pass for products with a boiling point of 370° C. or more into products with a boiling point of less than 370° C. was 85%.
The effluent from the hydroisomerization/hydrocracking step underwent separation, in a gas/liquid separator, of the unreacted hydrogen and light gases which were recycled to the hydroisomerization/hydrocracking step in order to recover a liquid effluent. The carbon monoxide (CO) content generated per pass in the gaseous effluent was limited to 1.1% by weight.
The liquid effluent from the separation step h) was then sent to a distillation train to separate the light products formed during these steps: the gases (C1-C4), a gasoline cut, a gas oil cut and a kerosene cut, and also a fraction, termed the residual fraction, which had an initial boiling point equal to 370° C. which was recycled in its entirety to the inlet to the hydroisomerization/hydrocracking reactor in order to maximize the production of gas oil and kerosene.
The yields are given in Table 6.
A process for producing middle distillates from a paraffinic effluent derived from a Fischer-Tropsch unit and divided into two fractions, a heavy fraction termed the waxes and a light fraction termed the cold condensate which was the same as that used in Example 1 was carried out, comprising a step for hydrogenation followed by a hydroisomerization/hydrocracking step, with no prior step for passage over at least one ion exchange resin; this was carried out for comparison purposes.
The light fraction, termed the cold condensate, was fractionated into a gaseous C4− fraction boiling at a temperature of less than 15° C. and an intermediate fraction with an initial boiling point in the range 15° C. to 40° C. and an end point in the range 300° C. to 400° C. The characteristics of the various fractions are given in Table 1 of Example 1.
The heavy fraction, termed the waxes, derived from the Fischer-Tropsch synthesis unit (line 11) passed over a guard bed composed of ACT 961 sold by Axens, at a temperature of 80° C., a pressure of 1 MPa, an hourly space velocity of 1 h−1. At the outlet, the effluent (line 13) contained less than 1 ppm of fines.
Said intermediate fraction which had not been purified by passage over at least one ion exchange resin and the wax fraction derived from the above decontamination step were mixed in their entirety. The composition of the mixture formed thereby is given in Table 7:
The recombined C5+ paraffinic fraction underwent a step for hydrogenation in the presence of hydrogen and a hydrogenation catalyst with trade name LD265 sold by Axens, said catalyst comprising 0.3% by weight of palladium deposited on an alumina with a specific surface area of 69 m2/g.
In order to maintain a conversion into olefins of 85% by weight, as for Example 1, the hydrogenation step was carried out at a reaction temperature of 150° C., at a pressure of 3.5 Mpa, the hydrogen was introduced at a flow rate such that the hydrogen/hydrocarbon volume ratio was 32 Nl/l/h and at an hourly space velocity of 8 h−1.
Under these conditions, the conversion into olefins was maintained at 85% by weight. The liquid effluent derived from the hydrogenation step had the composition described in Table 8:
All of the effluent from the hydrogenation step underwent a hydroisomerization/hydrocracking step in the presence of fresh hydrogen and a hydroisomerization/hydrocracking catalyst, in which the residual fraction with an initial boiling point of more than 370° C., unreacted hydrogen and light gases were recycled.
The hydroisomerization/hydrocracking catalyst comprised 0.3% by weight of platinum and a support comprising 29.3% by weight of silica, SiO2, and 70.7% by weight of alumina, Al2O3, a Na content of 10.0 ppm by weight, a total pore volume comprising 0.69 ml/g measured by mercury porosimetry, a volume of mesopores with a mean diameter of 80 Å representing 78% of the total pore volume measured by mercury porosimetry, a volume of macropores with a diameter of more than 500 Å representing 22% of the total pore volume measured by mercury porosimetry and a specific surface area of 300 m2/g.
The hydroisomerization/hydrocracking step was carried out under the conditions described in Table 9.
To maintain the conversion per pass for products with a boiling point of 370° C. or more into products with a boiling point of less than 370° C. at 85%, the temperature was increased and adjusted to 370° C.
The effluent from the hydroisomerization/hydrocracking step underwent separation, in a gas/liquid separator, of the unreacted hydrogen and light gases which were recycled to the hydroisomerization/hydrocracking step in order to recover a liquid effluent. The carbon monoxide (CO) content generated per pass in the gaseous effluent was limited to 2% by weight.
The presence of carbon monoxide (CO) derived from the decomposition of oxygen-containing compounds in the hydrocracking section not eliminated by passage over at least one ion exchange resin prior to hydrogenation and to hydroisomerization/hydrocracking, and being an inhibitor of the activity of the hydroisomerization/hydrocracking catalyst activity, necessitated an increase in the reaction temperature in order to maintain the conversion.
The liquid effluent from the separation step was then sent to a distillation train to separate the light products formed during these steps: the (C1-C4) gases, a gasoline cut, a gas oil cut and a kerosene cut, and also a fraction, termed the residual fraction, which had an initial boiling point equal to 370° C. which was recycled in its entirety to the inlet to the hydroisomerization/hydrocracking reactor in order to maximize the production of gas oil and kerosene.
The yields are given in Table 10.
The presence of carbon monoxide (CO) in the gaseous effluent which was an inhibitor of the hydrogenating function of the hydrocracking catalyst modified not only the activity but also the selectivity of said catalyst for middle distillates.
It can be seen that the absence of a prior step for passage over at least one ion exchange resin, necessitating an increase in temperature in order to maintain the conversion, resulted in an increase in the production of an unwanted light gas and naphtha fraction by over-cracking.
Thus, by using an ion exchange resin upstream of the hydrotreatment and hydroisomerization/hydrocracking steps, the process of the invention exemplified in Example 1 can reduce the total oxygen content of the feed and thereby limit the formation of carbon monoxide (CO) originating from the decomposition of oxygen-containing compounds present in the feed in the hydroisomerization/hydrocracking section. In fact, carbon monoxide (CO) is an inhibitor of the metallic compounds present on the hydroisomerization/hydrocracking catalyst and its content must be minimized in order not to require an increase in temperature in order to compensate for the drop in activity and to maintain conversion.
Thus, it can be seen that the process of the invention can reduce the production of carbon monoxide (CO) (1.1% by weight) by carrying out step b) of the invention compared with a process which is not in accordance with the invention which does not carry out said step for passage over at least one ion exchange resin, and can reduce the temperatures employed in the hydrogenation and hydroisomerization/hydrocracking steps in order to obtain the same conversion of 85% of products with a boiling point of 370° C. or more into products with a boiling point of less than 370° C.
| Number | Date | Country | Kind |
|---|---|---|---|
| 0901652 | Apr 2009 | FR | national |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/FR2010/000252 | 3/24/2010 | WO | 00 | 11/18/2011 |
