Patent Application
20150166908
The present invention relates to a process for intense desulphurization of a gasoline comprising diolefins, olefins and sulphur-containing compounds, including mercaptans, while minimizing the consumption of hydrogen and preserving the octane number.
The production of reformulated gasolines complying with new environmental standards in particular requires that their concentration of olefins be reduced slightly, but that their concentration of aromatics (in particular benzene) and sulphur be substantially reduced. Catalytically cracked gasolines, which may represent 30% to 50% of the gasoline pool, have high olefins and sulphur contents. Close to 90% of the sulphur present in reformulated gasolines can be attributed to gasoline from catalytic cracking (FCC, Fluid Catalytic Cracking). Desulphurization (hydrodesulphurization) of gasolines, principally FCC gasolines, is thus of substantial importance when complying with the specifications.
Pre-treatment by hydrotreatment (hydrodesulphurization) of feeds sent for catalytic cracking results in FCC gasolines typically containing less than 100 ppm of sulphur. These hydrotreatment units, however, operate under severe temperature and pressure conditions, which presupposes a high hydrogen consumption and high costs. In addition, the whole of the feed has to be desulphurized, which involves the treatment of very large volumes of feed.
Thus, in order to satisfy specifications as regards sulphur, it is necessary to post-treat the catalytically cracked gasolines by hydrotreatment (or hydrodesulphurization). When this post-treatment is carried out under conventional conditions known to the skilled person, it is possible to further reduce the sulphur content of the gasoline. However, this process suffers from the major disadvantage of causing a very large drop in the octane number for the FCC gasoline, due to saturation of olefins during the hydrotreatment.
U.S. Pat. No. 4,131,537 discloses the advantage of fractionating the gasoline into several cuts, preferably three, as a function of their boiling point, and of desulphurizing them under conditions which may be different and in the presence of a catalyst comprising at least one metal from group VIb and/or group VIII. That patent points out that the greatest benefit is obtained when the gasoline is fractionated into three cuts, and when the cut with intermediate boiling points is treated under mild conditions.
Patent FR 2 785 908 discloses the advantage of fractionating the gasoline into a light fraction and a heavy fraction then carrying out a specific hydrotreatment of the light gasoline over a nickel-based catalyst and a hydrotreatment of the heavy gasoline over a catalyst comprising at least one metal from group VIII and/or at least one metal from group VIb.
Of the possible pathways to producing fuels with a low sulphur content, that which has been very widely used consists of specifically treating the sulphur-rich gasoline bases by hydrodesulphurization processes in the presence of hydrogen. The traditional processes desulphurize the gasolines in a non-selective manner by hydrogenating a large proportion of the mono-olefins, which causes a large drop in the octane number and high hydrogen consumption. The most recent processes, such as the Prime G+ process (trade mark), can be used to desulphurize cracked gasolines which are rich in olefins while limiting the hydrogenation of mono-olefins and as a consequence the drop in the octane number. Processes of this type are described, for example, in patent applications EP 1 077 247 and EP 1 174 485.
As is disclosed in patent application EP 1 077 247, before the hydrotreatment step, it is advantageous to carry out a step for selective hydrogenation of the feed to be treated. This first hydrogenation step essentially consists of selectively hydrogenating the dienic compounds (diolefins), transforming the saturated light sulphur-containing compounds by weighting (by increasing their molecular weight), i.e. sulphur-containing compounds with a boiling point below the boiling point of thiophene, such as methanethiol, ethanethiol, as this means that simple distillation can then be applied to produce a light desulphurized gasoline fraction composed of a large quantity of olefins with no drop in the octane number.
The step for hydrodesulphurization of cracked gasolines which contain mono-olefins consists of passing the feed to be treated, mixed with hydrogen over a sulphide type transition metal catalyst in order to convert the sulphur-containing compounds into hydrogen sulphide (H2S). The reaction mixture is then cooled in order to condense the gasoline. The gas phase containing the excess hydrogen and H2S is separated and the desulphurized gasoline is recovered.
The residual sulphur-containing compounds generally present in the desulphurized gasoline may be separated into two distinct families: the non-hydrogenated sulphur-containing compounds present in the feed on the one hand, and the sulphur-containing compounds formed in the reactor by secondary reactions known as recombination reactions. In this latter family of sulphur-containing compounds, the major compounds are mercaptans obtained from the addition of the H2S formed in the reactor to the mono-olefins present in the feed. Mercaptans with chemical formula R—SH where R is an alkyl group are also known as recombination mercaptans and generally represent between 20% by weight and 80% by weight of the residual sulphur in the desulphurized gasolines.
The selective hydrogenation step described in patent application EP 1 077 247 is indispensable in order to prevent progressive deactivation of the selective hydrodesulphurization catalyst, to prevent gradual blocking of the reactor by the formation of polymerization gums on the catalyst surface or in the reactor, and to prevent the heat exchangers from clogging too quickly. Saturation of the diolefins is not generally necessary for the final application of the desulphurized gasoline. This selective hydrogenation step, which is indispensable only to the correct operation of the process in patent application EP 1 077 247, suffers from the disadvantage of inducing superfluous consumption of hydrogen linked to saturation of the diolefins of the feed. Since the diolefin hydrogenation is carried out under severe conditions, it is generally accompanied by slight hydrogenation of the olefins, which further increases the hydrogen consumption and induces a drop in octane number. Finally, in addition to the hydrogen consumed, hydrogen is also lost from the head of the distillation located between the selective hydrogenation step and the selective hydrodesulphurization step of the process of patent EP 1 077 247, since an excess of hydrogen is generally required in order to convert almost all of the diolefins in the first step.
U.S. Pat. No. 6,984,312 is also known, which discloses a process for the treatment of a light catalytically cracked gasoline (approximately C5-175° C.) containing olefins, diolefins, mercaptans and heavy sulphur-containing organic compounds. The process uses a first step for thioetherification in which the mercaptans are reacted with the diolefins in the feed in the presence of a thioetherification catalyst to form sulphides. The gasoline which has undergone this first step is then sent to a distillation column where it is fractionated into a light fraction which is depleted in sulphur and a heavy fraction containing the sulphides formed in the first step as well as heavy sulphur-containing organic compounds initially present in the gasoline to be treated. The heavy fraction is then treated in the presence of hydrogen and heavy cracked naphtha in a reactive distillation zone containing a hydrodesulphurization catalyst. The heavy cracked naphtha is recycled to the reactive distillation zone so that the distillation column can be operated at high temperature while retaining a liquid fraction in the catalytic bed.
The process described in U.S. Pat. No. 6,984,312 can obtain, with difficulty, a light fraction of gasoline which complies with future sulphur specifications, i.e. with an upper limit for total sulphur in the gasoline of 50 ppm (by weight), or even 30 or 10 ppm by weight (without post-treatment of this light fraction). In fact, the mercaptans are eliminated by addition to the diolefins of the feed with a thioetherification catalyst based on nickel or palladium. However, that type of catalyst also catalyses the selective hydrogenation of diolefins. Thus, the two reactions are concurrent on the diolefins, resulting in limited conversion of the mercaptans by thioetherification. The light gasoline produced from the head of the distillation column in the process described in U.S. Pat. No. 6,984,312 may thus contain a large fraction of light mercaptans present in the initial feed. For this reason, a post-treatment of the light fraction by hydrodesulphurization is necessary in order to further reduce the sulphur content of the light gasoline fraction.
The process described in U.S. Pat. No. 6,984,312 also suffers from the disadvantage of only being capable of treating a light gasoline and not a total gasoline. It has in fact been shown that a total gasoline (with a boiling point range which generally extends from 20° C. to 230° C.) cannot be treated effectively in a thioetherification reactor because of the large quantities of sulphur, which is poisonous to the thioetherification catalysts, in particular nickel-based catalysts (see U.S. Pat. No. 7,638,041) or catalysts based on palladium (see patent U.S. Pat. No. 5,595,634).
Thus, one aim of the invention is to propose a process for the production of a gasoline from a broad spectrum of gasoline types in terms of the boiling temperature range, with a low sulphur content, i.e. with a sulphur content of less than 50 ppm by weight and preferably less than 30 ppm or 10 ppm by weight, while limiting the consumption of hydrogen and the drop in octane number.
To this end, a process for the treatment of a gasoline comprising diolefins, olefins and sulphur-containing compounds including mercaptans is proposed, said process comprising the following steps:
a) carrying out a step for demercaptanization by addition of at least a portion of the mercaptans onto the olefins by bringing gasoline into contact with at least one first catalyst, at a temperature in the range 50° C. to 250° C., at a pressure in the range 0.4 to 5 MPa and with a LHSV in the range 0.5 to 10 h−1, the first catalyst being in the sulphide form and comprising a first support, at least one metal selected from group VIII and at least one metal selected from group VIb of the periodic classification of the elements, the % by weight, expressed in oxide equivalents of the metal selected from group VIII with respect to the total catalyst weight, is in the range 1% to 30% by weight and the % by weight, expressed in oxide equivalents of the metal selected from group VIb, is in the range 1% to 30% with respect to the total catalyst weight;
b) in a distillation column, carrying out fractionation of the gasoline obtained from step a) into at least one first intermediate light gasoline cut with a total sulphur content below that of the starting gasoline and a second intermediate heavy gasoline cut containing the major proportion of the starting sulphur-containing compounds;
c) introducing a stream of hydrogen and at least the second intermediate heavy gasoline cut obtained from step b) into a catalytic distillation column comprising at least one reaction zone including at least one second catalyst in the sulphide form comprising a second support, at least one metal from group VIII and a metal from group VIb, the conditions in the catalytic distillation column being selected so as to bring the intermediate heavy gasoline obtained from step b) into contact, in the presence of hydrogen, with the second catalyst in order to decompose the sulphur-containing compounds into H2S;
The process of the invention employs a first step a) in which the sulphur-containing mercaptan type compounds (R—SH) are transformed into heavier sulphur-containing compounds by reaction with the olefins present in the gasoline to be treated. The demercaptanization reactions of the invention are characterized by elimination of the mercaptans over the olefins:
This first step of weighting of the mercaptans reaches very high conversions (>90% and very often >95%), as the demercaptanization reactions occur selectively on the olefins which are generally present in high quantities. The lightest mercaptans are the most reactive in this step a).
Further, H2S, if it is present in the feed, is converted into mercaptan (which can itself be converted) by addition onto the olefins by means of a catalyst under selected conditions. This means that the presence of H2S is avoided in the overhead gases from step b), which contain mainly hydrogen which has not reacted in step a) and which it is advantageous to recycle to step a). This recycling, rendered possible by the absence of H2S in these gases, means that the hydrogen consumption for step a) can be still further reduced; this is an advantage of the process of the invention.
The demercaptanization reactions are preferably carried out over a catalyst comprising at least one metal from group VIII (groups 8, 9 and 10 of the new periodic classification of the elements, Handbook of Chemistry and Physics, 76th edition, 1995-1996), at least one metal from group VIb (group 6 of the new periodic classification of the elements, Handbook of Chemistry and Physics, 76th edition, 1995-1996) and a support.
Before bringing it into contact with the feed to be treated, the catalyst undergoes a sulphurization step. The catalyst only brings about the desired demercaptanization reaction when it is in its sulphide form. Sulphurization is preferably carried out in a sulphoreducing medium, i.e. in the presence of H2S and hydrogen, in order to transform the metallic oxides into sulphides such as MoS2 and Ni3S2.
In accordance with the invention, the process comprises a step b) for fractionation of the effluent obtained from demercaptanization step a) which is carried out in a fractionation column (or splitter). The column is configured so as to fractionate the gasoline into at least two cuts, namely: an intermediate light gasoline cut having a total sulphur content below that of the starting gasoline, and an intermediate heavy gasoline cut containing the major portion of the starting sulphur-containing compounds as well as the sulphur-containing products generated in step a). The distillation column of step b) may also be configured to function as a depentanizer or a dehexanizer. Preferably, the cut point for the distillation column of step b) is selected so as to avoid entraining thiophene in the intermediate light gasoline cut. The sulphur-containing compounds with a lower boiling point than that of thiophene (i.e. 84° C.) are converted in demercaptanization step a) and are thus entrained towards the intermediate heavy cut; the combination of steps a) and b) can be used to obtain an intermediate light gasoline cut with a very low sulphur content.
The intermediate light gasoline cut generally has a total sulphur content of less than 50 ppm by weight, preferably less than 30 ppm or even less than 10 ppm and contains at least all of the C5 olefins, preferably the C5 compounds and at least 20% by weight of the C6 olefins. Recovery of a large proportion of the olefins of the feed in the intermediate light gasoline cut means that the selectivity of the process as regards the hydrogenation of olefins is significantly improved and overconsumption of hydrogen can be avoided, because the olefins are not redirected towards the selective hydrodesulphurization section and thus will not run any risk of being hydrogenated.
The intermediate heavy gasoline cut generally contains hydrocarbons having, for example, a boiling point of more than 84° C., the heavy sulphur-containing compounds (from the thiophene, sulphide, disulphide families) initially present in the gasoline to be treated, as well as the sulphur-containing compounds which are essentially of the sulphide type which are formed during step a) by addition of mercaptans onto the olefins.
In accordance with the invention, the process comprises a step c) for hydrodesulphurization of the heavy fraction obtained from the fractionation step b). This treatment is carried out in a distillation column into which a catalytic reaction zone, also known as a “catalytic column”, has been incorporated.
This step c) consists of desulphurizing the feed in the catalytic column by contact with hydrogen injected into the column and a hydrodesulphurization catalyst.
The distillation column is thus configured to operate under conditions which can allow the sulphur-containing compounds of the feed (mercaptans, sulphides and thiophene compounds) to react with hydrogen to form H2S.
At the same time that step c) takes place in the catalytic distillation column, a step d) takes place in which separation of the feed constituted by at least the second intermediate heavy fraction obtained from the fractionation step b), optionally mixed with a recycle stream, into at least two fractions is carried out, namely into a desulphurized final light gasoline fraction originating from the decomposition of sulphur-containing compounds, and into a desulphurized heavy fraction.
The final light gasoline cut is recovered from the head of the catalytic column with the H2S produced by desulphurization and unreacted hydrogen, while the desulphurized heavy fraction is generally evacuated from the lower portion or even from the bottom of the catalytic column.
Optionally, a complementary desulphurized gasoline cut may be withdrawn as a side stream from a point located between the inlet and the bottom of the column.
The final light gasoline cut, accompanied by H2S and hydrogen which has not reacted in the catalytic column, is then condensed in order to separate the noncondensables from the liquid phase. A portion of this final light gasoline cut is withdrawn, while the other is recycled to the column as an internal reflux.
Optionally, the process of the invention also comprises a step for recycling all or a portion of the desulphurized heavy gasoline cut to the catalytic distillation column. When this optional recycle is used, a makeup and a purge of the recycle stream may also be carried out.
In accordance with an embodiment in which recycling of the desulphurized heavy gasoline cut is total and there is no complementary withdrawal, the fractionation operated by the catalytic column acts only to recirculate the desulphurized heavy gasoline cut. In this case, the final light gasoline cut withdrawn from the head of the catalytic distillation column is constituted by the second intermediate heavy gasoline cut obtained from step b). The desulphurized heavy cut which is withdrawn from the bottom of the column is constituted in this case by a hydrocarbon cut added by means of a makeup and with a boiling point in a temperature range which is higher than the second intermediate heavy gasoline cut obtained from step b). This external hydrocarbon cut, which is then recycled as a loop to the catalytic distillation column, acts to maintain a liquid phase at the bottom of the catalytic column in order to operate the column at a higher temperature in order to desulphurize the heaviest sulphur-containing molecules which are also the most difficult to convert. The bottom of the catalytic column operates at high temperature and is the zone of the catalytic bed which is most sensitive to deactivation by the deposition of coke or gums. This heavier cut is preferably of the paraffinic type and acts as a solvent to wash the coke and the gums which are deposited at the bottom of the catalytic column. This washing is indispensable in order to obtain very good cycle times for the catalytic bed. This is all the more true when the feed for the catalytic column has a large quantity of unsaturated diolefin type compounds.
This embodiment is particularly advantageous for limiting the consumption of hydrogen in step a), because a step for hydrogenation of the gum and coke precursors (in particular diolefin type compounds) is no longer necessary.
In accordance with another embodiment in which there is no recycling of the desulphurized heavy gasoline cut, the intermediate heavy gasoline cut obtained from fractionation step b) is separated into at least two cuts, each of these cuts being desulphurized. In this configuration, the two cuts recovered from the outlet from the catalytic column can be upgraded directly for the gasoline pool.
In accordance with another embodiment, the catalytic distillation column is operated so as to separate the second intermediate heavy gasoline cut obtained from the fractionation step b) into at least two desulphurized cuts. This embodiment uses a makeup with an external heavy hydrocarbon cut, inter alia. This makeup of hydrocarbons is recycled to said catalytic distillation column in order to maintain a liquid phase at the bottom of the catalytic column. In this case, the two desulphurized cuts obtained from the second intermediate heavy gasoline cut are respectively withdrawn from the head (final light fraction) and by means of a side stream (complementary desulphurized gasoline cut), and the desulphurized heavy cut which is withdrawn from the bottom constitutes the recycled heavy cut.
In the context of the invention, it is also possible to carry out partial recycles.
One advantage of the process of the invention resides in the fact that it is not necessary to desulphurize the light fraction of the gasoline obtained from fractionation step b) because almost all of the mercaptan type sulphur-containing compounds have been transformed into compounds with a higher molecular weight during step a), and so they are entrained in the heavy gasoline fraction. This gasoline fraction has a low sulphur content and a good octane number and does not need post-treatment.
The hydrogenation reactions are not required in step a). Hydrogen, if it is used, essentially acts to maintain a hydrogenating surface condition for the catalyst so as to ensure a high yield for the demercaptanization reactions. The process of the invention is thus not penalized by the low pressures and entails a reduced consumption of hydrogen, which is an advantage of the process of the invention.
Another advantage of the process is that the first two steps can be carried out at the same pressure (apart from the pressure drop) as step a) only requires a little hydrogen, or even none at all, which is also the case with step b). The absence of a need for the dedienization reaction during step a) is also favourable in terms of the hydrogen consumption, as little or no hydrogen is consumed during this step. This iso-pressure operation for steps a) and b) means that the gas at the head of the column for step b), which is rich in hydrogen, can be recycled towards the demercaptanization reactor of step a) when the catalyst for step a) has to have a hydrogenating surface condition which is appropriate for high demercaptanization conversions. This recycling means that the hydrogen consumption in step a) can be reduced, and thus the loss of this hydrogen to the fuel gas network can be prevented. This hydrogen normally contains no H2S, as it is not produced by the catalyst used in step a) under the selected conditions. This H2S may even be converted in step a) if it is present in the feed.
One advantage of the process of the invention is based on the fact that, in contrast to the thioetherification reactors described in the prior art, the catalyst and the operating conditions used during step a) can treat a whole gasoline (i.e. C5-220° C.) with a high sulphur content. Treating a whole gasoline cut using this catalyst is particularly advantageous in maintaining a liquid phase in step c) when high desulphurization conversions are envisaged, i.e. when the catalytic column functions at high temperature.
The use of a catalytic column and not a conventional fixed bed hydrodesulphurization reactor allows for continuous rinsing of the catalytic zone by the liquid reflux inside the column. This rinsing of the catalytic zone means that coking of the catalyst can be reduced, and thus the cycle time for the hydrodesulphurization catalyst of step c) can be extended. This rinsing of the catalytic zone also means that gums, which may be formed by diolefin polymerization, can be washed. The partial pressure of hydrogen is also reduced compared with conventional fixed bed hydrodesulphurization, which is favourable to preventing side reactions of olefin hydrogenation, which generates both an overconsumption of hydrogen and a drop in octane number.
Another advantage in relation to the use of a catalytic column for carrying out selective hydrodesulphurization is based on the fact that the continuous rising stream of hydrogen can entrain the H2S produced by the hydrodesulphurization reactions and thus contribute to limiting the formation of recombination mercaptans by the addition of hydrogen sulphide onto the olefins which are still present.
The aim of the present invention is to provide a process for the desulphurization of a gasoline having a limited sulphur content starting from a gasoline preferably obtained from a catalytic cracking, coking or visbreaking unit. The gasoline may be a “total” cracked gasoline (C5-220° C.) or a gasoline with a final boiling point of 210° C. or less (light gasoline).
In accordance with the invention, the gasoline first undergoes a step a) for transformation of the sulphur-containing compounds, essentially the lightest mercaptans of the gasoline, over olefins in order to increase their molecular weight. The process also comprises a second step b) which consists of passing all or a portion of the gasoline obtained in step a) into a fractionation column also known as a “splitter”.
This concatenation can be used to obtain a light fraction the sulphur content of which has been reduced without substantial reduction of the olefins content, even for intense levels of desulphurization, and without the need to treat this light gasoline using a supplemental hydrodesulphurization section or having recourse to processes which can restore the octane number of the gasoline.
Thus, the process of the invention can be used to provide a light gasoline fraction which can be sent directly to the gasoline pool, which has a total sulphur content of less than 50 ppm, preferably less than 30 ppm, or even less than 10 ppm, depending on the quantity of sulphur initially present and the chemical nature of the sulphur-containing compounds.
The process of the invention also comprises a step c) for hydrodesulphurization of the heavy fraction obtained from the fractionation step b). This treatment is carried out in a distillation column into which a catalytic reaction zone has been incorporated, also termed the “catalytic column”.
This second step consists of desulphurizing this first heavy fraction by contact with hydrogen over the catalytic bed.
The catalytic distillation column is configured to operate under conditions which can simultaneously:
In the context of the present application, the expression “catalytic column” designates an apparatus in which the catalytic reaction and separation of the products takes place at least simultaneously. The apparatus employed may comprise a distillation column equipped with a catalytic section in which the catalytic reaction and distillation take place simultaneously. It may also be a distillation column in association with at least one reactor disposed inside said column and on a wall thereof. The internal reactor may be operated as a vapour phase reactor or as a liquid phase reactor with circulation of the liquid/vapour as a co-current or as a counter-current.
Using a catalytic distillation column has the advantage over the use of a single fixed bed reactor operated in the gas phase of allowing continuous rinsing of the catalytic zone by the reflux liquid inside the column. This rinsing of the catalytic zone means that coking of the catalyst can be reduced, and also the cycle time of the hydrodesulphurization catalyst can be extended. The partial pressure of hydrogen is also reduced compared with conventional fixed bed hydrodesulphurization, which is favourable as regards preventing side reactions of olefin hydrogenation which generate both an overconsumption of hydrogen and a drop in octane number. The use of a catalytic column also means that the reaction can be controlled, while favouring exchange of the heat released; the heat of reaction can be absorbed by the heat of evaporation of the mixture.
The process of the invention can be used to treat any type of gasoline cut containing sulphur, preferably a gasoline cut obtained from a catalytic cracking unit, for which the boiling point range typically extends from approximately the boiling points of hydrocarbons containing 2 or 3 carbon atoms (C2 or C3) to approximately 250° C., more preferably from approximately the boiling points of hydrocarbons containing 5 carbon atoms to approximately 220° C.
Thus, the process of the invention is also applicable to a gasoline cut which has already been stabilized, i.e. a gasoline cut from which the hydrocarbons containing fewer than 6 or 5 carbon atoms has been removed.
The process of the invention can also be used to treat a gasoline feed termed “light” with a final boiling point of less than those mentioned above such as, for example, 210° C. or less, 180° C. or less, 160° C. or less or 145° C. or less.
The sulphur content of the gasoline cuts produced by catalytic cracking (FCC) depends on the sulphur content of the feed treated by FCC, the presence or otherwise of a pre-treatment of the FCC feed, as well as the end point of the cut. In general, the sulphur contents of the whole of a gasoline cut, in particular those from FCC, are more than 100 ppm by weight and most of the time more than 500 ppm by weight. For gasolines with end points of more than 200° C., the sulphur contents are often more than 1000 ppm by weight, and may in some cases even reach values of the order of 4000 to 5000 ppm by weight.
As an example, the gasolines obtained from catalytic cracking units (FCC) contain, on average, between 0.5% and 5% by weight of diolefins, between 20% and 50% by weight of olefins, between 10 ppm and 0.5% by weight of sulphur, generally including less than 300 ppm of mercaptans. The mercaptans are generally concentrated in the light fractions of the gasoline, and more precisely in the fraction with a boiling point below 120° C.
It should be noted that the sulphur-containing compounds present in the gasoline may also comprise heterocyclic sulphur-containing compounds such as thiophenes, alkylthiophene or benzothiophenes, for example.
Step a) for Weighting the Mercaptans with Olefins
This step consists of transforming light sulphur-containing compounds from the mercaptans family, i.e. compounds which are in the light gasoline after the fractionation step b), into heavier sulphur-containing compounds which are entrained in the intermediate heavy gasoline fraction during fractionation step b).
During this step a), a demercaptanization reaction occurs which consists in addition of the mercaptans to the olefins of the feed in the presence of a catalyst.
Typically, the mercaptans which can react during step a) are as follows (non-exhaustive list): methyl mercaptan, ethyl mercaptan, n-propyl mercaptan, isopropyl mercaptan, isobutyl mercaptan, tert-butyl mercaptan, n-butyl mercaptan, sec-butyl mercaptan, isoamyl mercaptan, n-amyl mercaptan, α-methylbutyl mercaptan, α-ethyl propyl mercaptan, n-hexyl mercaptan and 2-mercaptohexane.
The demercaptanization reaction is preferably carried out over a catalyst comprising at least one metal from group VIII (groups 8, 9 and 10 of the new periodic classification of the elements, Handbook of Chemistry and Physics, 76th edition, 1995-1996), at least one metal from group VIb (group 6 of the new periodic classification of the elements, Handbook of Chemistry and Physics, 76th edition, 1995-1996) and a support. The metal from group VIII is preferably selected from nickel and cobalt, and in particular nickel. The metal from group VIb is preferably selected from molybdenum and tungsten and highly preferably is molybdenum.
The support for the catalyst is preferably selected from alumina, nickel aluminate, silica, silicon carbide, or a mixture of these oxides. Preferably, alumina is used and more preferably, pure alumina. Preferably, a support with a total pore volume, measured by mercury porosimetry, in the range 0.4 to 1.4 cm3/g is used, preferably in the range 0.5 to 1.3 cm3/g. The specific surface area of the support is preferably in the range 70 m2/g to 350 m2/g. In a preferred variation, the support is a cubic gamma alumina or delta alumina.
The catalyst used in step a) generally comprises:
In particular, it has been found that the performance of the catalysts is improved when the catalyst has the following characteristics:
In a preferred embodiment of the invention, a catalyst is used containing in the range 4% to 12% by weight of nickel oxide (in the form NiO), in the range 6% to 18% by weight of molybdenum oxide (in the form MoO3) and with a nickel/molybdenum molar ratio in the range 1 to 2.5, the metals being deposited on a support constituted solely by alumina and with a degree of sulphurization of the metals constituting the catalyst of more than 80%.
The catalyst of the invention may be prepared using any technique which is known to the skilled person, in particular by impregnation of the metals from groups VIII and VIb onto the selected support.
After introducing the metals from groups VIII and VIb, and optional shaping of the catalyst, it undergoes an activation treatment. This treatment is generally intended to transform the molecular precursors of the metals into the oxide phase. In this case it is an oxidizing treatment, but simple drying of the catalyst may also be carried out. In the case of an oxidizing treatment, also known as calcining, this is generally carried out in air or in diluted oxygen, and the treatment temperature is generally in the range 200° C. to 550° C., preferably in the range 300° C. to 500° C.
After calcining, the metals deposited on the support are in the oxide form. In the case of nickel and molybdenum, the metals are principally in the form of MoO3 and NiO. Before contact with the feed to be treated, the catalysts undergo a sulphurization step. Sulphurization is preferably carried out in a sulphoreducing medium, i.e. in the presence of H2S and hydrogen, in order to transform the metallic oxides into sulphides such as, for example, MoS2 and Ni3S2. Sulphurization is carried out by injecting a stream containing H2S and hydrogen over the catalyst, or a sulphur-containing compound which is capable of decomposing into H2S in the presence of the catalyst and hydrogen. Polysulphides such as dimethyldisulphide are H2S precursors which are in routine use for catalyst sulphurization. The temperature is adjusted so that the H2S reacts with the metallic oxides to form metallic sulphides. This sulphurization may be carried out in situ or ex situ (inside or outside the reactor) as regards the demercaptanization reactor, at temperatures in the range 200° C. to 600° C. and more preferably in the range 300° C. to 500° C.
Step a) may be carried out without adding hydrogen to the reactor, but preferably it is injected with the feed so as to maintain a hydrogenating surface condition for the catalyst which is appropriate for high levels of demercaptanization conversion. Typically, step a) functions with a H2 flow rate/feed flow rate ratio in the range 0 to 25 Nm3 of hydrogen per m3 of feed, preferably in the range 0 to 10 Nm3 of hydrogen per m3 of feed, highly preferably in the range 0 to 5 Nm3 of hydrogen per m3 of feed, and more preferably in the range 0.5 to 2 Nm3 of hydrogen per m3 of feed.
The whole of the feed is generally injected into the reactor inlet. However, it may be advantageous in certain cases to inject a fraction or all of the feed between two consecutive catalytic beds placed in the reactor. This embodiment in particular means that the reactor can continue to be operated if the reactor inlet becomes blocked by deposits of polymers, particles or gums present in the feed.
The gasoline to be treated is brought into contact with the catalyst at a temperature in the range 50° C. to 250° C., preferably in the range 80° C. to 220° C., and more preferably in the range 90° C. to 200° C., with a liquid hourly space velocity (LHSV) in the range 0.5 h−1 to 10 h−1, the liquid hourly space velocity unit being the litre of feed per litre of catalyst per hour (L/L·h). The pressure is in the range 0.4 MPa to 5 MPa, preferably in the range 0.6 to 2 MPa, and more preferably in the range 0.6 to 1 MPa.
At the end of step a), the gasoline treated under the conditions listed above has a reduced mercaptans content. In general, the gasoline produced contains less than 50 ppm by weight of mercaptans, preferably less than 10 ppm by weight. More than 80%, or even more than 90% of the light sulphur-containing compounds with a boiling point less than that of thiophene (84° C.) is generally converted. The olefins are not or are only slightly hydrogenated, which means that a good octane number can be maintained at the outlet from step a). The degree of olefin hydrogenation is generally less than 2%.
Step b) for Separation into an Intermediate Light Gasoline Cut and an Intermediate Heavy Gasoline Cut
Separation step b) is preferably carried out by means of a conventional distillation column also known as a “splitter”. This fractionation column can be used to separate an intermediate light gasoline fraction containing a small amount of sulphur-containing compounds and an intermediate heavy gasoline fraction preferably containing the major portion of the sulphur-containing compounds initially present in the initial gasoline.
This column generally operates at a pressure in the range 0.1 to 2 MPa, preferably in the range 0.6 to 1 MPa. It should be noted that this pressure may be substantially the same (with the exception of the pressure drop) as that prevailing in the reactor for step a). This iso-pressure operation of steps a) and b) means that the overhead gas from the column of step b), which is rich in hydrogen, can be recycled to the demercaptanization reactor of step a) (when the catalyst for step a) has to have a hydrogenating surface condition appropriate to high demercaptanization conversions). This recycle means that the hydrogen consumption in step a) can be reduced and means that loss of this hydrogen to the fuel gas network can be prevented. This hydrogen normally does not contain H2S as it is not produced by the catalyst used in step a) under the selected conditions.
The number of theoretical plates in this separation column is generally in the range 10 to 100, preferably in the range 20 to 60. The reflux ratio, expressed as the ratio of the liquid flow rate in the column divided by the flow rate of distillate expressed in kg/h, is generally less than 1, preferably less than 0.8.
The intermediate light gasoline obtained at the end of separation b) generally contains at least all of the C5 olefins, preferably the C5 compounds and at least 20% of the C6 olefins. The cut point for the column is often determined so that thiophene is not entrained in the intermediate light gasoline cut. Thus, the intermediate heavy gasoline cut has an initial point located around 84° C. This initial point may optionally be higher, depending on the anticipated sulphur content in the intermediate light gasoline, and may be from approximately 100° C. to 120° C.
Alternatively, the distillation column is configured to allow withdrawal of an intermediate gasoline cut as a side stream, i.e. a gasoline cut with boiling points included between the final boiling point of the intermediate light gasoline and the initial boiling point of the intermediate heavy gasoline. Said intermediate gasoline may then be treated by hydrodesulphurization in a dedicated reactor and then mixed with the intermediate light gasoline.
The reaction for desulphurization of step c) is a hydrodesulphurization reaction carried out by passing the feed, in the presence of hydrogen which is injected into said column, over at least one catalyst which is at least partially in the sulphide form, comprising at least one metal from group VIII, at least one metal from group VIb and optionally phosphorus, at a temperature in the range 210° C. to 350° C., preferably in the range 220° C. to 320° C. The pressure at the column head is generally maintained at between approximately 0.1 and approximately 4 MPa, preferably in the range 1 to 3 MPa. The H2 flow rate/feed flow rate ratio in the column is in the range 25 to 400 Nm3 per m3 of liquid feed, preferably in the range 40 to 100 Nm3 per m3 of liquid feed.
The metal from group VIII is preferably cobalt or nickel, and the metal from group VIb is generally molybdenum or tungsten. Combinations such as cobalt-molybdenum or nickel-molybdenum are preferred. The quantity of metal from group VIII, expressed as the oxide, is generally in the range 0.5% to 25% by weight, preferably in the range 1% to 10% by weight with respect to the weight of catalyst. The quantity of metal from group VIb, expressed as the oxide, is generally in the range 1.5% to 60% by weight, preferably in the range 3% to 50% by weight with respect to the weight of catalyst.
Preferably, when the catalyst is of the cobalt-molybdenum type, the quantity of cobalt, expressed as the oxide, is generally in the range 0.5% to 15% by weight, and more preferably in the range 2% to 5% by weight; the quantity of molybdenum, expressed as the oxide, is in the range 1.5% to 60% by weight, and more preferably in the range 5% to 20% by weight.
Preferably, when the catalyst is of the nickel-molybdenum type, the quantity of nickel, expressed as the oxide, is generally in the range 0.5% to 25% by weight, more preferably in the range 5% to 25% by weight; the quantity of molybdenum, expressed as the oxide, is in the range 1.5% to 30% by weight, more preferably in the range 3% to 20% by weight.
The support for the catalyst is usually a porous solid such as, for example, an alumina, a silica-alumina or other porous solids such as, for example, magnesia, silica or titanium oxide, used alone or as a mixture with alumina or silica-alumina, and may start out in the form of extrudates with a small diameter, or as spheres. The catalyst in the column must have a structural shape which is suitable for catalytic distillation in order to act both as a catalytic agent for carrying out the reactions and also as a material transfer agent in order to provide separation stages throughout the length of the bed.
In order to minimize hydrogenation of the olefins of the treated feed, it is advantageous and preferable to use a cobalt-molybdenum type catalyst in which the density of molybdenum, expressed as the % by weight of MoO3 per unit surface area, is more than 0.07 and preferably more than 0.12. The catalyst of the invention preferably has a specific surface area of less than 250 m2/g, more preferably less than 230 m2/g, and highly preferably less than 190 m2/g.
If a good hydrodesulphurization conversion is to be obtained at the same time as olefin hydrogenation (in particular at the bottom of the catalytic column when a heavy gasoline cut is recirculated), it is advantageous and preferable to use a nickel-molybdenum type catalyst. The catalyst of the invention in this case preferably has a specific surface area in the range 70 to 250 m2/g.
The metals are deposited on the support using any methods known to the skilled person such as, for example, dry impregnation, or excess impregnation of a solution containing the precursors of the metals. Said solution is selected so as to be able to dissolve the precursors of the metals in the desired concentrations. In the case of synthesis of a CoMo catalyst, for example, the molybdenum precursor may be molybdenum oxide, or ammonium heptamolybdate. Examples for cobalt which may be cited are cobalt nitrate, cobalt hydroxide and cobalt carbonate. The precursors are generally dissolved in a medium which allows them to be dissolved in the desired concentrations. Thus, depending on the case, it may be carried out in an aqueous medium and/or in an organic medium. The phosphorus may be added in the form of phosphoric acid.
In the context of the invention, it is possible to use more than one catalytic bed in the reaction zone, for example two distinct catalytic beds, separated from each other via a gap. It is also possible in certain configurations for the catalytic column to comprise more than one catalytic bed charged with different catalysts, in particular when an effective combination of different properties of the cobalt-molybdenum and nickel-molybdenum type catalysts is to be employed. In another configuration of the process of the invention, the catalytic bed of the column may be located solely above the infeed or solely below it. Preferably, the column has one or more catalytic beds covering at least a portion of both the zone located above the infeed and the zone located below the infeed.
Operation of the catalytic column induces the simultaneous presence of vapour and liquid in the reaction zone. A large portion of the vapour is constituted by hydrogen, the remainder being constituted by a portion of the vaporized feed and hydrogen sulphide.
As is the case with any distillation, there is a temperature gradient in the system such that the lower end of the column comprises compounds with a boiling point which is higher than that of the upper end of the column. The distillation can be used to separate the compounds present in the feed by boiling point difference.
The heat of reaction which may be generated in the catalytic column is evacuated by vaporizing the mixture on the distillation plate concerned. As a consequence, the thermal profile of the column is very stable and the catalytic reactions which occur on the bed do not perturb its operation. Similarly, this stability of the thermal profile means that stable reaction kinetics are obtained as they are isothermal on each separation stage, the temperatures being dependent only on the liquid-vapour equilibrium of the separation stages and on the control of the pressure in the column.
The catalytic distillation column is configured so as to be able to function under operating conditions which can be used to separate the feed constituted by at least the second intermediate heavy fraction obtained from the fractionation step b) into at least two fractions, namely a desulphurized final light gasoline fraction deriving from the decomposition of the sulphur-containing compounds, and a desulphurized heavy fraction.
The final light gasoline cut is recovered at the head of the catalytic column with the H2S produced by desulphurization and the unreacted hydrogen, while the desulphurized heavy fraction is withdrawn from the bottom of the catalytic column.
Optionally, a complementary desulphurized gasoline cut can be withdrawn as a side stream at a point located between the infeed and the column bottom.
Preferably, the final light gasoline cut accompanied by the H2S produced by the desulphurization reactions and unreacted hydrogen are cooled to a temperature which is generally below 60° C. in order to condense the hydrocarbons. The gas phases (containing mainly the H2S produced and unreacted hydrogen) and the liquid hydrocarbon phase (i.e. the upgradable final light gasoline cut) are separated in a separator. A portion of this final light gasoline cut is transferred to the gasoline pool, while another is recycled to the column as an internal reflux. The internal reflux is both useful for carrying out distillation of the feed and also can be used as a permanent wash for the catalyst. The downflow of the liquid in the column means that the catalyst can be cleaned of the coke and gums which may be formed, primarily due to the presence of highly unsaturated compounds, of the diolefin or acetylene type in the feed. This means that deactivation of the catalyst can be reduced and as a consequence the cycle time is improved.
The H2S and hydrogen-rich gas phase may, for example, be sent to an amines absorber in order to purify and recover hydrogen with a view to recycling it to the process.
Optionally, the process of the invention also comprises a step for recycling, to the catalytic column, all or a portion of the desulphurized heavy gasoline cut which is withdrawn from the bottom of said catalytic column. When this optional recycle is used, a makeup and purge of the recycle stream may also be employed.
Alternatively, the recycle stream may also comprise an external hydrocarbon cut (supplied via the makeup) with an initial boiling point which is greater than or equal to that of the intermediate heavy gasoline cut. This external hydrocarbon cut is withdrawn from the bottom of the catalytic distillation column and recycled to said column in a loop.
Various configurations of the process of the invention are indicated below; the list of these configurations is not exhaustive.
In a first configuration, the feed entering step c), which is constituted by the intermediate heavy gasoline cut obtained from step b), is fractionated into at least two gasoline cuts. In this case, the final light gasoline cut and the desulphurized heavy cut obtained from step c) and d) can be upgraded directly to the gasoline pool. This configuration is preferably used when the gasoline cut injected into step a) at the inlet to the entire concatenation is a total gasoline, i.e. where the boiling point range typically extends from approximately the boiling points of hydrocarbons containing 2 or 3 carbon atoms (C2 or C3) to approximately 250° C. or, as is preferable, from approximately the boiling points of hydrocarbons containing 2 or 3 carbon atoms (C2 or C3) to approximately 220° C. or, as is more preferable, from approximately the boiling points of hydrocarbons containing 5 carbon atoms to approximately 220° C. Nevertheless, it may also be carried out with a light gasoline, i.e. with a final boiling point of less than 210° C.
In this configuration, the catalytic bed is preferably composed of a single bed of cobalt-molybdenum type catalyst.
The final light gasoline cut recovered from the head and the desulphurized heavy cut recovered from the bottom are gasoline cuts which have been desulphurized to a very low sulphur content, i.e. with a sulphur content of less than 50 ppm by weight, preferably less than 30 ppm or 10 ppm by weight. The final light gasoline cut is generally a cut with a boiling point range which typically extends from approximately the initial point of the feed for the catalytic column, generally in the range 80° C. to 120° C., up to approximately 145° C. or, as is preferable, to approximately 160° C. or, as is more preferable, to approximately 180° C. The desulphurized heavy cut is generally a cut with a boiling point range which typically extends from approximately the end point of the final light gasoline cut to approximately the end point of the feed for the catalytic column, generally in the range 220° C. to 250° C. The two cuts recovered from the outlet from the catalytic column may then be mixed and sent to a stripper in order to eliminate the last traces of dissolved H2S so that it can finally be sent to the gasoline pool.
In a second configuration, the feed for step c) comprises the intermediate heavy gasoline cut obtained from step b) and a recycle of all or a portion of a desulphurized heavy cut recovered from the bottom of the catalytic column. This configuration is particularly used when the gasoline cut injected into step a) at the inlet to the whole of the concatenation is a light gasoline with a final boiling point of less than 220° C. such as, for example, 210° C. or less, 180° C. or less, 160° C. or less or indeed 145° C. or less.
However, when the intermediate heavy gasoline cut recovered from the bottom of the distillation column of step b) has a final boiling point of approximately 180° C., it is desirable to use a makeup of an external heavy gasoline cut as a recycle in order to maintain a liquid phase in the catalytic column under the selected operating conditions. This recycle can also be used to increase the rate of washing the catalyst at the bottom, which is favourable when the feed for the hydrodesulphurization step contains coke and gum precursors.
Preferably, a single makeup of the heavy cut is made into the recycle stream and said heavy cut is recycled to the column as a loop. This external recycled gasoline cut typically has a distillation range from the end point of the feed to be treated, i.e. in a range from 145° C. to approximately 210° C. for this configuration, to a temperature in the range from approximately 180° C. to 240° C. This external recycled gasoline cut may, for example, be a desulphurized cracked heavy gasoline cut. The external recycled heavy cut must have a low unsaturated compound content so that it can be used as a solvent for optimized washing of the catalyst.
In this configuration, the catalytic column will preferably contain two catalytic beds respectively located above and below the inlet. Preferably, a catalyst is loaded into the bottom of the catalytic column which has both hydrodesulphurization and hydrogenation properties. The catalyst comprises at least one metal from group VIII and at least one metal from group VIb in the sulphide form; preferably, the metal from group VIII is nickel and the metal from group VIb is molybdenum. In contrast, the catalyst located in the upper zone is preferably a cobalt-molybdenum type catalyst.
In a third configuration, three different cuts are withdrawn from the catalytic column:
This configuration is preferably used when the gasoline cut injected into step a) at the inlet to the whole concatenation is a total gasoline, i.e. for which the boiling point range typically extends from approximately the boiling points of hydrocarbons containing 2 or 3 carbon atoms (C2 or C3) to approximately 250° C. or, as is preferable, from approximately the boiling points of hydrocarbons containing 2 or 3 carbon atoms (C2 or C3) to approximately 220° C. or, more preferably, from approximately the boiling points of hydrocarbons containing 5 carbon atoms to approximately 220° C. This configuration is also preferred when the catalytic column has to operate at very high conversions (intense desulphurization), and thus at high temperatures, in particular when the end point for the cut treated in the process of the invention is particularly high and the feed thus contains heavy sulphur-containing compounds, of the thiophene or even benzothiophene type, which are difficult to desulphurize.
Recycling the heavy cut to the catalytic column means that, despite the high temperature, a liquid phase can be maintained in the column and also means that the flow rate for washing the catalyst at the bottom can be increased. In fact, operating at a higher temperature favours the formation of coke and gums by polymerization of the diolefins in the feed, in particular at the column bottom where the temperatures are the highest.
Preferably, a makeup of an external hydrocarbon cut is added to the recycle loop. This external heavy cut typically has a distillation range from 220° C. to 270° C., preferably 220° C. to 250° C. This heavy cut is generally a cracked heavy cut obtained from fractionation of FCC such as a LCO (Light Cycle Oil, i.e. a cut obtained from catalytic cracking and boiling in a temperature range higher than that of gasoline) or a kerosene cut or a straight run diesel.
In this configuration, the catalytic column will preferably contain two catalytic beds respectively located above and below the infeed. The catalyst for the zone located at the bottom of the catalytic column will preferably be a catalyst of the nickel-molybdenum type. The catalyst located in the upper zone is preferably a catalyst of the cobalt-molybdenum type which will provide for good selectivity of the hydrodesulphurization reactions compared with those for hydrogenation of the olefins in order to maintain the octane number of the treated feed.
The final light gasoline cut recovered from the head and the complementary desulphurized gasoline cut recovered as a side stream are desulphurized gasoline cuts with a low sulphur content, i.e. with a sulphur content of below 50 ppm by weight, preferably less than 30 ppm or 10 ppm by weight. The final light gasoline cut is generally a cut with a boiling point range which typically extends from approximately the initial point for the feed treated in the catalytic column (generally in the range 80° C. to 120° C.) to approximately a temperature of more than 145° C., or preferably to approximately a temperature of more than 160° C., or more preferably to approximately 180° C. The complementary desulphurized gasoline cut is generally a cut with a boiling point range which typically extends from approximately the end point for the final light gasoline cut to approximately the end point for the second intermediate heavy gasoline cut obtained from step b), i.e. to approximately a temperature in the range 210° C. to 230° C. The two gasoline cuts recovered from the catalytic column outlet (final light gasoline and complementary desulphurized gasoline) may then be mixed then sent to a stripper in order to eliminate the last traces of dissolved H2S so that they can finally be stored in the gasoline pool.
These and other aspects of the invention will be clarified in the detailed description of particular embodiments of the invention, made with reference to the figures in the drawings, in which:
In general, similar elements are designated by identical reference numerals in the figures.
In accordance with the process, the gasoline feed to be treated is sent with an optional makeup of hydrogen to a demercaptanization reactor 2 by means of a feed line 1.
The reactor 2 comprises a catalytic section provided with a catalytic bed specifically selected to carry out selective addition of mercaptans to the olefins with a view to increasing their molecular weight.
The reactor is preferably a fixed catalytic bed reactor which operates in a three-phase or two-phase system with one of the phases (the catalyst) being solid.
The demercaptanization reactions are generally carried out at a temperature in the range 50° C. to 250° C., at a pressure in the range 0.6 to 2 MPa and at a liquid hourly space velocity in the range 0.5 h−1 to 10 h−1.
The effluent obtained from demercaptanization step a) is then sent to a fractionation column 4 which is also known as a “splitter”, via the line 3. The fractionation column 4 is configured and operated so as to separate an intermediate light gasoline cut containing a low sulphur fraction and an intermediate heavy gasoline cut containing the major portion of the sulphur initially present in the gasoline to be treated. This column generally operates at a pressure in the range 0.1 to 2 MPa, preferably in the range 0.6 to 1 MPa. The number of theoretical plates for this fractionation column is generally in the range 10 to 100, preferably in the range 20 to 60. The reflux ratio, expressed as the ratio of the liquid passing through the column divided by the flow rate of distillate expressed in kg/h, is generally less than 1 and preferably less than 0.8. The intermediate light gasoline obtained from the separation generally contains at least all of the C5 olefins, preferably the C5 compounds and at least 20% of the C6 olefins. In general, this light fraction has a very low sulphur content, i.e. less than 50 ppm by weight, preferably less than 30 ppm by weight, or even less than 10 ppm by weight. It is not necessary to post-treat the light cut before using it as a gasoline base.
As shown in
The intermediate heavy gasoline cut which is withdrawn from the bottom of the fractionation column 4 and which contains the major portion of the sulphur-containing compounds including those generated during the demercaptanization step a) acts as a feed for the third step of the process of the invention.
Referring now to
In order to carry out catalytic conversion of the sulphur-containing compounds, hydrogen is supplied via the line 16.
The catalytic distillation column 14 is configured so as to carry out fractionation of said intermediate heavy gasoline into at least two fractions, namely a desulphurized final heavy gasoline cut and a desulphurized final light gasoline cut. The two cuts, the desulphurized final light gasoline and the desulphurized final heavy gasoline, may then be sent to a stripper in order to eliminate the final traces of dissolved H2S (not shown).
As can be seen in
The final light gasoline distilling at the head of the column accompanied by the H2S formed following the desulphurization reactions and hydrogen which has not reacted in the column is then cooled by means of a heat exchanger 18 then sent via the line 19 to a gas/liquid separator 20 where a gaseous effluent essentially comprising hydrogen and H2S are separated (via the line 21) along with a desulphurized liquid gasoline. The desulphurized liquid gasoline is then divided into two fractions, one fraction being recycled to the distillation column 14 in order to provide a reflux and another fraction which may be used in a gasoline pool after optionally passing through a H2S stripper. The overhead gases may be sent to an amines absorption unit in order to separate hydrogen from hydrogen sulphide to purify the hydrogen with a view to possible recycling. The process of the invention as shown in
A second embodiment of the process of the invention is shown in
A third embodiment of the process of the invention is shown in
A fourth embodiment of the process is shown in
Preferably, the hydrodesulphurization catalyst comprises at least one support, at least one metal from group VIII (groups 8, 9 and 10 of the new periodic classification of the elements, Handbook of Chemistry and Physics, 76th edition, 1995-1996) and at least one metal from group VIb (group 6 of the new periodic classification of the elements, Handbook of Chemistry and Physics, 76th edition, 1995-1996). Preferably, the catalyst has a density of metals from group VIb per unit surface area of the support in the range (limits included) 2×10−4 to 18×10−4 g of oxides of metals from group VIb per m2 of support, preferably in the range (limits included) 3×10−4 to 16×10−4 g of oxides of metals from group VIb per m2 of support, more preferably in the range (limits included) 3×10−4 to 14×10−4 g of oxides of metals from group VIb per m2 of support, and very preferably in the range (limits included) 4×10−4 to 13×10−4 g of oxides of metals from group VIb per m2 of support.
The quantity, expressed with respect to the total catalyst weight, of metals from group VIb is preferably in the range (limits included) 1% to 20% by weight of oxides of metals from group VIb, more preferably in the range (limits included) 1.5% to 18% by weight of oxides of metals from group VIb, highly preferably in the range (limits included) 2% to 15% by weight of oxides of metals from group VIb, and still more preferably in the range (limits included) 2.5% to 12% by weight of oxides of metals from group VIb. Preferably, the metal from group VIb is molybdenum or tungsten or a mixture of these two metals, and more preferably, the metal from group VIb is constituted solely by molybdenum or tungsten. Highly preferably, the metal from group VIb is molybdenum.
The quantity, expressed with respect to the total catalyst weight, of metals from group VIII is preferably in the range (limits included) 0.1% to 20% by weight of oxides of metals from group VIII, more preferably in the range (limits included) 0.2% to 10% by weight of oxides of metals from group VIII, more preferably in the range (limits included) 0.3% to 5% by weight of oxides of metals from group VIII. Preferably, the metal from group VIII is cobalt or nickel or a mixture of these two metals, and more preferably, the metal from group VIII is constituted solely by cobalt or nickel. Highly preferably, the metal from group VIII is cobalt.
The molar ratio of the metals from group VIII to metals from group VIb is generally in the range (limits included) 0.1 to 0.8, preferably (limits included) in the range 0.2 to 0.6, and more preferably (limits included) in the range 0.3 to 0.5.
The hydrodesulphurization catalyst may further comprise phosphorus. The phosphorus content is preferably in the range (limits included) 0.1% to 10% by weight of P2O5, more preferably (limits included) in the range 0.2% to 5% by weight of P2O5, highly preferably (limits included) in the range 0.3% to 4% by weight of P2O5, still more preferably (limits included) in the range 0.35% to 3% by weight of P2O5, with respect to the total catalyst weight.
When phosphorus is present, the molar ratio of phosphorus to the metal from group VIb is generally 0.25 or more, preferably 0.27 or more, more preferably in the range (limits included) 0.27 to 2, still more preferably in the range (limits included) 0.35 to 1.40, highly preferably in the range (limits included) 0.45 to 1.10, and still more preferably in the range (limits included) 0.45 to 1.0, or even (limits included) in the range 0.50 to 0.95.
The support for the catalyst is a porous solid selected from the group constituted by: aluminas, silica, silica-aluminas or even oxides of titanium or magnesium, used alone or as a mixture with alumina or silica-alumina. It is preferably selected from the group constituted by: silica, the transition alumina family and silica-aluminas; highly preferably, the support is essentially constituted by at least one transition alumina, i.e. it comprises at least 51% by weight, preferably at least 60% by weight, highly preferably at least 80% by weight, or even at least 90% by weight of transition alumina. It may optionally be constituted solely by a transition alumina.
The specific surface area of the support used to prepare the catalyst before incorporation of the metals from groups VIb and VIII, optionally shaped and heat treated, is generally less than 200 m2/g, preferably less than 170 m2/g, more preferably less than 150 m2/g, highly preferably less than 135 m2/g, or even less than 100 m2/g and even less than 85 m2/g. The support may be prepared using any precursor, any preparation method and any shaping tool known to the skilled person.
1000 cc of a NiMo 8/8 catalyst on a nickel aluminate support in the form of 2-4 mm spheres was charged into a fixed bed downflow reactor. It was initially sulphurized by injecting a feed of heptane containing 4% DMDS at a hydrogen flow rate of 500 N litre/litre over 4 h, at a HSV of 2 h−1, at 350° C. and at 2.5 MPa. Under these conditions, the DMDS decomposed to form H2S and allowed sulphurization of the catalyst to take place.
The feed used for the test was a FCC gasoline with an initial boiling point IP=2° C. and a final boiling point FP=208° C.
The operating conditions were as follows:
An analysis by sulphur-containing compound species provided the following:
It will be seen that under these conditions, a conversion of 97.6% was obtained for the C1 to C3 mercaptans. These mercaptans are the sulphur-containing compounds which are the most susceptible of being found in the light fraction of the gasoline after distillation.
Chromatographic analysis of the feed and the effluent provided the following results for the hydrocarbon families:
It will be seen that the olefins were almost unhydrogenated between the inlet and the outlet of the reactor. Thus, the octane number was not degraded.
In addition, the chromatographic method used allowed the C5 diolefins to be identified, which were withdrawn along with the olefins family. These diolefins were: isoprene, 1,3-cis-pentadiene and 1,3-trans-pentadiene. Their conversion was approximately 17% in the reactor.
Finally, the measurement of the MAV (Maleic Anhydride Value) provided us with information regarding the quantity of highly unsaturated compounds present in the outlet effluents. It was observed that the reactor effluents had a MAV very close to the feed.
The effluent obtained from the reactor was then separated in a batch mode distillation column. The effluent was charged into a 100 L reboiler heated by resistances, while the condensation was ensured at the column head by water supplemented with glycol in order to prevent the loss of light compounds. The water supplemented with glycol was at a temperature of 15° C.
The column had a diameter of 10 cm and was filled with packing (multiknit pads) over a height of 2 m. Separation was carried out with a reflux ratio of 15. The separation pressure was atmospheric pressure. When the head thermocouple reached a temperature of 65° C. and the bottom temperature was about 90° C., distillation was stopped. The target cut point was 65° C.
This batch distillation meant that pilot scale separation could be carried out, reproducing the industrial splitter with the following characteristics:
The light gasoline fraction recovered from the column head represented 32.8% by weight of the initial gasoline.
The characteristics of the products recovered overhead (after stripping the H2S formed) and from the bottom were as follows:
The combination of the demercaptanization reactor with a splitter thus meant that an intermediate light gasoline cut with a very low sulphur content could be recovered.
The heavy gasoline cut obtained from the bottom of the splitter was then sent to the catalytic distillation column. This heavy gasoline cut, termed the intermediate cut, had a MAV which was a little higher than the feed due to the separation.
The intermediate heavy gasoline cut was injected into a catalytic distillation column with a diameter of 5 cm and a height of 12 m.
This column was charged with 0.75 kg of hydrodesulphurization catalyst based on cobalt and molybdenum supported on an alumina in the sulphide form. This catalyst contained 3% by weight of cobalt in the oxide form and 10% by weight of molybdenum in the oxide form. The feed was injected in the presence of hydrogen such that 70% by weight of the catalyst was below the level of the infeed. The catalytic distillation column functioned under the following operating conditions:
The results for the overhead (after degassing H2S) and bottom cuts were as follows:
| Number | Date | Country | Kind |
|---|---|---|---|
| 1202028 | Jul 2012 | FR | national |
| 1352290 | Mar 2013 | FR | national |
| Filing Document | Filing Date | Country | Kind |
|---|---|---|---|
| PCT/FR2013/051418 | 6/18/2013 | WO | 00 |
