Patent Application
20080306316
The present invention relates to a process for converting an aromatic hydrocarbon which comprises aromatic sulfur compounds, or a mixture of aromatic hydrocarbons which comprises aromatic sulfur compounds, if appropriate in the presence of hydrogen, wherein, in a first step, aromatic sulfur compounds are removed (step a), and, in a second step, the aromatic hydrocarbon or the mixture of aromatic hydrocarbons is hydrogenated in the presence of a supported ruthenium catalyst in the presence of hydrogen (step b).
In one embodiment, the present invention relates to a process in which the aromatic hydrocarbon is benzene. In a further embodiment, the present invention relates to a process wherein a mixture of aromatic hydrocarbons is used. In this case, it is possible, for example, to use mixtures which comprise benzene and toluene. However, it is also possible to use mixtures which comprise benzene and xylene or a xylene isomer mixture, or mixtures which comprise benzene, toluene and xylene or a xylene isomer mixture. In step a), the content of aromatic sulfur compounds, for example thiophene, is lowered to ≦70 ppb, and the total sulfur content to a total of ≦200 ppb, and, in step b), the desulfurized aromatic hydrocarbon or the desulfurized mixture of aromatic hydrocarbons is reduced in the presence of a supported ruthenium catalyst and hydrogen to the corresponding cycloaliphatic hydrocarbon or the corresponding mixture of corresponding cycloaliphatic hydrocarbons. In the case of benzene, the hydrogenation product obtained is thus cyclohexane, that obtained from toluene is methylcyclohexane and that obtained from xylene is the dimethylcyclohexane corresponding in each case, and that obtained from a xylene isomer mixture is the corresponding dimethylcyclohexane isomer mixture which can be purified by distillation.
There exist numerous processes for hydrogenating benzene to cyclohexane. These hydrogenations are carried out predominantly over nickel and platinum catalysts in the liquid or gas phase (here, cf., inter alia, U.S. Pat. No. 3,597,489, U.S. Pat. No. 2,898,387, GB 799,396). Typically, the majority of the benzene is first hydrogenated to cyclohexane in a first reactor and then the unconverted amount of benzene is converted to cyclohexane in one or more downstream reactors.
The strongly exothermic hydrogenation reaction requires careful temperature and residence time control in order to achieve full conversion at high selectivity. In particular, significant formation of methylcyclopentane, which proceeds preferentially at relatively high temperatures, has to be suppressed. Typical cyclohexane specifications require a residual benzene content of <100 ppm and a methylcyclopentane content of <200 ppm. The content of n-paraffins (for example n-pentane, n-hexane) is likewise also critical. These undesired compounds are likewise formed preferentially at relatively high hydrogenation temperatures and, just like methylcyclopentane, can be removed from the desired cyclohexane only by complicated separating operations (for example extraction, rectification or use of molecular sieves, as described in GB 1,341,057). The catalyst used too has a strong influence on the degree of formation of undesired secondary components, such as methylcyclohexane, n-hexane, n-pentane, etc.
In view of this background, it is desirable to carry out the hydrogenation at minimum temperatures. On the other hand, this is limited, since, depending on the type of hydrogenation catalyst used, an adequately high hydrogenation activity of the catalyst, which is in turn sufficient for an economically viable space-time yield, is achieved only from relatively high temperatures.
The nickel and platinum catalysts used for the benzene hydrogenation additionally have a series of disadvantages, Nickel catalysts are very sensitive toward sulfur-containing impurities in benzene, so that either very pure benzene has to be used for the hydrogenation, or, as described in GB 1,104,275, a platinum catalyst which tolerates a higher sulfur content is used in the main reactor, thus protecting the postreactor which comprises a nickel catalyst.
Another possibility is to dope the hydrogenation catalyst with rhenium, as described in GB 1,155,539, or to incorporate ion exchangers into the hydrogenation catalyst, as disclosed in GB 1,144,499. However, the preparation of such catalysts is complicated and expensive.
Platinum catalysts have fewer disadvantages than nickel catalysts, but are very expensive.
As an alternative, the recent literature has therefore referred to ruthenium-containing catalysts for hydrogenating benzene to cyclohexane.
SU 319 582 describes ruthenium suspension catalysts which have been doped with palladium, platinum or rhodium for preparing cyclohexane from benzene. However, these are very expensive owing to the palladium, platinum or rhodium used, and the workup and recovery of the catalyst is additionally both complicated and expensive in the case of suspension catalysts.
U.S. Pat. No. 3,917,540 describes Al2O3-supported catalysts for preparing cyclohexane from benzene. As the active metal, these comprise a noble metal from transition group VIII of the Periodic Table, and also an alkali metal and technetium or rhenium. Also described in U.S. Pat. No. 3,244,644 are η-Al2O3-supported ruthenium hydrogenation catalysts which are also said to be suitable for hydrogenating benzene. However, these catalysts comprise at least 5% active metal. Moreover, the preparation of η-Al2O3 is both complicated and expensive.
In addition, WO 00/63142 describes, inter alia, the hydrogenation of unsubstituted aromatics using a catalyst which comprises, as the active metal, at least one metal of transition group VIII of the Periodic Table and which has been applied to a support having macropores. Suitable active metals are in particular ruthenium and suitable supports are in particular appropriate aluminum oxides and zirconium dioxides.
One advantage of these processes lies in the comparatively favorable costs of ruthenium which is used as the active metal for the catalyst in comparison to the costs which arise as a result of other hydrogenation metals such as palladium, platinum or rhodium. However, a disadvantage here too is that these ruthenium catalysts are sensitive toward sulfur impurities.
EP 600 406 discloses that unsaturated hydrocarbons such as alkenes (for example ethene) which are contaminated with thiophene can be desulfurized by treating the unsaturated hydrocarbon in the presence of a copper-zinc desulfurizing agent which has a copper/zinc atomic ratio of 1:about 0.3-10, and which is obtainable by a co-precipitation process, with from 0.01 to 4% by volume of hydrogen. In particular, it is emphasized that the amount of hydrogen should not exceed these values, since this leads to undesired hydrogenation of the unsaturated hydrocarbons to be purified.
It was a primary object of the present invention to provide a process for hydrogenating aromatic hydrocarbons or mixtures thereof which comprise aromatic sulfur compounds to the corresponding cycloaliphatics or mixtures thereof, in particular benzene to obtain cyclohexane, and which enables cycloaliphatics, or the mixtures thereof, to be obtained with very high selectivity and space-time yield.
Accordingly, the present invention relates to a process for converting an aromatic hydrocarbon which comprises aromatic sulfur compounds, or a mixture of aromatic hydrocarbons which comprises aromatic sulfur compounds, wherein, in a first step, aromatic sulfur compounds, if appropriate in the presence of hydrogen, are removed (step a); this desulfurization is carried out in the presence of a copper-zinc desulfurizing agent which has a copper:zinc atomic ratio of from 1:0.3 to 1:10 and is obtainable by a coprecipitation process. In a second step, the aromatic hydrocarbon thus obtained or the mixture of aromatic hydrocarbons thus obtained is hydrogenated in the presence of a supported ruthenium catalyst and hydrogen to give the corresponding cycloaliphatics or mixtures thereof (step b), the catalyst having been applied to a support which has meso- and/or macropores.
In a preferred embodiment, the aromatic hydrocarbon used is benzene which is hydrogenated to cyclohexane in the presence of hydrogen.
In a further preferred embodiment, a mixture of aromatic hydrocarbons is used, which is hydrogenated to the corresponding mixture of cycloaliphatics in the presence of hydrogen. Useful mixtures of aromatic hydrocarbons are those which comprise benzene and toluene, or benzene and xylene or a xylene isomer mixture, or benzene, toluene and xylene or a xylene isomer mixture. The hydrogenation affords cyclohexane from benzene, methylcyclohexane from toluene, and the corresponding dimethylcyclohexanes from the xylenes.
In step a), the aromatic hydrocarbon or the mixture of aromatic hydrocarbons, each of which comprises aromatic sulfur compounds as an impurity, is desulfurized. Possible aromatic sulfur-containing impurities are particularly thiophene, benzothiophene, dibenzothiophene or corresponding alkylated derivatives, in particular thiophene. In addition to these aromatic sulfur compounds, it is also possible for further sulfur-containing impurities, for example hydrogen sulfide, mercaptans such as methyl mercaptan, tetrahydrothiophene, disulfides such as dimethyl disulfide, COS or CS2, referred to hereinafter as nonaromatic sulfur compounds, to be present in the aromatic hydrocarbon or the mixture of aromatic hydrocarbons. In addition, other impurities may also be present, such as water, C5-C7-alkanes, for example n-heptane, C5-C7-alkenes, for example pentene or hexene, where the double bond may be present at any position in the carbon skeleton, C5-C7-cycloalkanes, for example methylcyclopentane, ethylcyclopentane, dimethylcyclopentane, cyclohexane, methylcyclohexane, or C5-C7 cycloalkenes, for exam pie cyclohexene.
The aromatic hydrocarbon used in a particular embodiment generally has a purity of >98% by weight, in particular >99% by weight, preferably >99.5% by weight, especially preferably >99.9% by weight. When a mixture of aromatic hydrocarbons is used, the fraction of aromatic hydrocarbons in the mixture used is >98% by weight, in particular >99% by weight, preferably >99.5% by weight, especially preferably >99.9% by weight. In both cases, the content of aromatic sulfur-containing impurities may be up to 2 ppm by weight, preferably up to 1 ppm by weight. The content of total sulfur impurities may be up to 5 ppm by weight, preferably up to 3 ppm by weight, in particular up to 2 ppm by weight, specifically up to 1 ppm by weight. Other impurities may be up to 2% by weight, preferably up to 0.5% by weight, in particular up to 0.10% by weight. Water may be present in the aromatic hydrocarbon or in the corresponding mixtures of aromatic hydrocarbons up to 0.1% by weight, preferably up to 0.07% by weight, in particular up to 0.05% by weight.
The desulfurization is carried out over a copper-zinc desulfurizing agent, if appropriate in the presence of hydrogen. This copper-zinc desulfurizing agent comprises at least copper and zinc, the copper, zinc atomic ratio being in the range from 1:0.3 to 1:10, preferably from 1:0.5 to 1:3 and in particular from 1:0.7 to 1:1.5. It is obtained by a coprecipitation process and can be used in oxidized or else in reduced form.
In a particular embodiment, the copper-zinc desulfurizing agent comprises at least copper, zinc and aluminum, the copper:zinc:aluminum atomic ratio being in the range from 1:0.3:0.05 to 1:10:2, preferably from 1:0.6:0.3 to 1:3:1 and in particular from 1:0.7:0.5 to 1:1.5:0.9.
The desulfurizing agent can be prepared by various processes. For example, an aqueous solution which comprises a copper compound, especially a water-soluble copper compound, for example copper nitrate or copper acetate, and a zinc compound, especially a water-soluble zinc compound, for example zinc nitrate or zinc acetate, together with an aqueous solution of an alkaline substance (for example sodium carbonate, potassium carbonate) can be mixed with one another to form a precipitate (coprecipitation process). The precipitate formed is filtered off, washed with water or first washed, then filtered and subsequently dried. Calcination is then effected at from about 270 to 400° C. Subsequently, the solid obtained is slurried in water, filtered off and dried. The copper-zinc desulfurizing agent thus obtained (“oxidized form”) can be used in the desulfurization in this form.
In a further embodiment, it is possible to subject the mixed oxide thus obtained to a hydrogen reduction. This is carried out at from about 150 to 350° C., preferably at from about 150 to 250° C., in the presence of hydrogen, the hydrogen being diluted by an inert gas, for example nitrogen, argon, methane, especially nitrogen, so that the hydrogen content is 10% by volume or less, preferably 6% by volume or less, in particular from 0.5 to 4% by volume. The copper-zinc desulfurizing agent thus obtained (“reduced form”) can be used in the desulfurization in this form.
In addition, the copper-zinc desulfurizing agent may also comprise metals which belong to group VIII of the Periodic Table (such as Fe, Co, Ni, Ru, Rh, Pd, Os, Ir, Pt), to group IB (such as Ag, Au) or to group VIB (such as Cr, Mo, W). These can be prepared by adding the appropriate metal salts to the abovementioned preparation processes.
It is also possible to shape or to extrude the solid obtained after the calcination or else that obtained after the hydrogen treatment to tablets or to other shapes, in which case it may be helpful to add additives, for example binders, for example graphite.
In a further embodiment, a solution which comprises a copper compound, especially a water-soluble copper compound, for example copper nitrate or copper acetate, a zinc compound, especially a water-soluble zinc compound, for example zinc nitrate or zinc acetate, and an aluminum compound, for example aluminum hydroxide, aluminum nitrate, sodium aluminate, together with an aqueous solution of an alkaline substance, for example sodium carbonate, potassium carbonate, can be mixed with one another to form a precipitate (coprecipitation process). The precipitate formed is filtered off, washed with water, or first washed, then filtered and dried. Calcination is then effected at from about 270 to 400° C. Subsequently, the solid obtained is slurried in water, filtered off and dried. The copper-zinc desulfurizing agent thus obtained (“oxidized form”) can be used in the desulfurization in this form.
In a further embodiment, it is possible to subject the mixed oxide thus obtained to a hydrogen reduction. This is carried out at from about 150 to 350° C., preferably at from about 150 to 250° C., in the presence of hydrogen, the hydrogen being diluted by an inert gas, for example nitrogen, argon, methane, in particular nitrogen, so that the hydrogen content is 10% by volume or less, preferably 6% by volume or less, in particular from 0.5 to 4% by volume. The copper-zinc desulfurizing agent thus obtained (“reduced form”) can be used in the desulfurization in this form.
In addition, the copper-zinc desulfurizing agent may also comprise metals which belong to group VII of the Periodic Table (such as Fe, Co, Ni, Ru, Rh, Pd, Os, Ir, Pt), to group IB (such as Ag, Au) or to group VIB (such as Cr, Mo, W). These can be prepared by adding the appropriate metal salts to the abovementioned preparation processes.
It is also possible to shape or to extrude the solid obtained after the calcination or else that obtained after the hydrogen treatment to tablets or to other shapes, in which case it may be helpful to add additives, for example binders, for example graphite.
In a further embodiment, the coprecipitation can be carried out under pH control, for example, by adjusting the feed rate of the salt solutions such that a pH of from about 7 to 7.5 is maintained during the precipitation. It is also possible to subject the precipitate which is formed in the precipitation, after washing, to spray-drying.
In a further embodiment, the coprecipitation can be carried out in such a way that the copper oxide-zinc oxide components are precipitated from aqueous solutions of the corresponding salts (for example nitrates or acetates) with an alkaline substance (for example alkali metal carbonate, ammonium carbonate) in the presence of aluminum oxide, aluminum hydroxide in colloidal distribution (as a gel or sol).
The calcination, the hydrogen treatment which may be desired and the shaping can be effected as described above.
It is also possible to use commercially available catalysts, for example the catalyst R 3-12 from BASF or G-132A from Süd-Chemie.
In a preferred embodiment, the copper-zinc desulfurizing agent is used in reduced form. It may be advantageous to subject the mixed oxide which is obtained by the above-described processes to a hydrogen reduction which can be carried out as follows ([cat] hereinafter represents catalyst):
The copper-zinc desulfurizing agent thus obtained is then present in the reduced form and can be used thus. However, it can also be stored under inert gas until it is used. In addition, it is also possible to store the copper-zinc desulfurizing agent in an inert solvent. From case to case, it may be advantageous to store the copper-zinc desulfurizing agent in its oxidized form and to carry out the activation just in time. In this connection, it may also be advantageous to carry out a drying step before the activation. In this case, the calcined copper-zinc desulfurizing agent present in oxidic form is heated in a nitrogen stream of from 200 to 400 m3 (STP)/m3[CAT]·h, in particular of 300±20 m3 (STP)/m3[CAT]·h, to from 180 to 220° C., in particular to 200±10° C., at a heating rate which should not exceed 50 K/h. As soon as the water has been removed, cooling can be effected to from 100 to 140° C., in particular to 120±5° C., at a cooling rate which should not exceed 50 K/h, and the activation can be carried out as described above.
In an especially preferred embodiment, a copper-zinc desulfurizing agent is used which comprises from 35 to 45% by weight, preferably from 38 to 41% by weight, of copper oxide, from 35 to 45% by weight, preferably 38 to 41% by weight, of zinc oxide, and from 10 to 30% by weight, preferably from 18 to 24% by weight, of aluminum oxide, and if appropriate further metal oxides.
In an exceptionally preferred embodiment, a copper-zinc desulfurizing agent is used which comprises from 38 to 41% by weight of copper oxide, from 38 to 41% by weight of zinc oxide, and from 18 to 24% by weight of aluminum oxide.
These copper-zinc desulfurizing agents are obtainable from corresponding calcined mixed oxides by the abovementioned preparation processes.
In one embodiment, the desulfurization of the aromatic hydrocarbon or of the mixture of aromatic hydrocarbons, preferably of benzene, is carried out over the copper-zinc desulfurizing agent in oxidized form without addition of hydrogen.
In one further embodiment, the desulfurization of the aromatic hydrocarbon or of the mixture of aromatic hydrocarbons, preferably of benzene, is carried out over the copper-zinc desulfurizing agent in oxidized form in the presence of hydrogen.
In one further embodiment, the desulfurization of the aromatic hydrocarbon or of the mixture of aromatic hydrocarbons, preferably of benzene, is carried out over the copper-zinc desulfurizing agent in reduced form without addition of hydrogen.
In one further embodiment, the desulfurization of the aromatic hydrocarbon or of the mixture of aromatic hydrocarbons, preferably of benzene, is carried out over the copper-zinc desulfurizing agent in reduced form in the presence of hydrogen.
Typically, the desulfurization is carried out within a temperature range of from 40 to 200° C., particularly at from 50 to 180° C., in particular at from 60 to 160° C., preferably at from 70 to 120° C., at a pressure of from 1 to 40 bar, particularly at from 1 to 32 bar, preferably at from 1.5 to 5 bar, in particular at from 2.0 to 4.5 bar. The desulfurization may be carried out in the presence of inert gases, for example nitrogen, argon or methane. In general, however, the desulfurization is carried out without addition of inert gases.
Typically, if desired, hydrogen is used here which has a purity of ≧99.8% by volume, in particular of ≧99.9% by volume, preferably of ≧99.95% by volume. These purities apply analogously to the hydrogen which is used in the activations of the catalysts carried out if appropriate.
Typically the weight ratio of aromatic hydrocarbon or of the mixture of aromatic hydrocarbons to hydrogen is in the range from 40 000:1 to 1000:1, particularly in the range from 38 000:1 to 5000:1, especially in the range from 37 000:1 to 15 000:1, preferably in the range from 36 000:1 to 25 000:1, specifically in the range from 35 000:1 to 30000:1.
In general, the LHSV (“Liquid Hourly Space Velocity”) is in the range from 0.5 to 10 kg of aromatic hydrocarbon per part by volume of catalyst and hour (kg/(m3[cat]·h)), in particular in the range from 1 to 8 kg/(m3[cat]·h), preferably in the range from 2 to 6 kg/(m3[cat]·h).
The aromatic hydrocarbon or the mixture of aromatic hydrocarbons, preferably benzene, thus desulfurized now has a content of aromatic sulfur compounds of at most 70 ppb, preferably at most 50 ppb, and the total sulfur content is a total of ≦200 ppb, preferably ≦150 ppb, in particular ≦100 ppb.
The above-described desulfurizing agents also enable chlorine, arsenic and/or phosphorus or corresponding chlorine, arsenic and/or phosphorus compounds to be reduced or to be removed from the aromatic hydrocarbon or the mixture of aromatic hydrocarbons.
The aromatic hydrocarbon or the mixture of aromatic hydrocarbons can be desulfurized in one or more reactors connected in parallel or in series. These reactors are typically operated in liquid-phase mode, the gas and the liquid being conducted in cocurrent or in countercurrent, preferably in cocurrent. However, the possibility also exists of operating the reactors in trickle mode, the gas and the liquid being conducted in cocurrent or in countercurrent, preferably in countercurrent.
If necessary, the desulfurizing agent can also be removed again from the reactor. When the desulfurizing agent is present in reduced form, it may be advantageous to subject the desulfurizing agent to an oxidation before the deinstallation. The oxidizing agents used are oxygen or mixtures of oxygen with one or more inert gases, for example air. The oxidation is effected by customary processes known to those skilled in the art. For example, the oxidation can be carried out as follows:
The copper-zinc desulfurizing agent thus obtained can then be deinstalled.
In step b), the desulfurized aromatic hydrocarbon or the mixture of aromatic hydrocarbons is then hydrogenated in the presence of a supported ruthenium catalyst to the corresponding cycloaliphatics or the corresponding mixtures of cycloaliphatics, the catalyst having been applied to a support which has meso- and/or macropores.
The supports used may in principle be all supports which have macropores, i.e. supports which have exclusively macropores and also those which also comprise mesopores and/or micropores in addition to macropores. The terms “macropores”, “mesopores” and “micropores” are used in the context of the present invention as defined in Pure Appl. Chem. 46, 71 (1976), specifically as pores whose diameter is above 50 nm (macropores) or whose diameter is between 2 and 50 nm (mesopores) or whose diameter is <2 nm (micropores).
Especially suitable as supports are appropriate activated carbons, silicon carbides, aluminum oxides, silicon oxides, titanium dioxides, zirconium dioxides, or else mixtures thereof. Preference is given to using appropriate aluminum oxides, zirconium dioxides or silicon oxides, especially γ-aluminum oxide or silicon oxides.
The active metal used may be ruthenium alone or together with at least one further metal of transition groups IB, VIIB or VII of the Periodic Table of the Elements (CAS version). Suitable further active metals in addition to ruthenium are, for example, platinum, rhodium, palladium, iridium, cobalt or nickel or a mixture of two or more thereof. Among the metals of transition groups IB and/or VIIB of the Periodic Table of the Elements which can likewise be used, suitable metals are, for example, copper and/or rhenium. Preference is given to using ruthenium alone as the active metal or together with platinum or iridium in the coated catalyst; very particular preference is given to using ruthenium alone as the active metal.
An essential constituent of the catalysts is the support material based on silicon dioxide, generally amorphous silicon dioxide. In this context, the term “amorphous” is understood to mean that the fraction of crystalline silicon dioxide phases makes up less than 10% by weight of the support material. However, the support materials used to prepare the catalysts may have superstructures which are formed by regular arrangement of pores in the support material.
The hydrogenation of the desulfurized aromatic hydrocarbon or of the mixture of desulfurized aromatic hydrocarbons, preferably benzene, over the above-described supported ruthenium catalysts to the cycloaliphatics or the corresponding mixture of cycloaliphatics, preferably cyclohexane, in the presence of hydrogen, can be carried out in the liquid phase or in the gas phase. The hydrogenation process is preferably carried out in the liquid phase—generally at a temperature of from 50 to 250° C., preferably at from 60 to 200° C., in particular at from 70 to 170° C. The pressures used are in the range from 1 to 200 bar, preferably from 10 to 50 bar, in particular from 19 to 40 bar and especially from 25 to 35 bar.
Typically hydrogen with a purity of ≧99.8% by volume, in particular of ≧99.9% by volume, preferably of ≧99.95% by volume, is used in the hydrogenation.
More preferably, the aromatic hydrocarbon or the mixture of aromatic hydrocarbons is hydrogenated fully, full hydrogenation being understood to mean a conversion of the compound to be hydrogenated of generally >98%, preferably >99%, more preferably >99.5%, even more preferably >99.9%, in particular >99.99% and especially >99.995%.
Typically, the weight ratio of aromatic hydrocarbon or of the mixture of aromatic hydrocarbons to hydrogen is in the range from 8:1 to 5:1, preferably from 7.7:1 to 5.5:1, in particular from 7.6:1 to 6:1 and especially from 7.5:1 to 6.5:1.
The hydrogenation of the desulfurized aromatic hydrocarbon or mixtures of desulfurized aromatic hydrocarbons can be carried out in one reactor or in a plurality of reactors connected in series or parallel, which are preferably operated in trickle mode. In this case, the gas and the liquid are conducted in cocurrent or in countercurrent, preferably in cocurrent. However, it is also possible to operate the reactors connected in series in liquid-phase mode.
In general, the LHSV (“Liquid Hourly Space Velocity”) is in the range from 0.1 to 10 kg of aromatic hydrocarbon per part by volume of catalyst and hour (kg/(m3[cat]·h)), preferably in the range from 0.3 to 1.5 kg/(m3[cat]·h). The trickle density is typically in the range from 20 to 100 m3 of aromatic hydrocarbon per unit of cross-sectional area of the catalyst bed available for flow and hour (m3/m2·h), preferably in the range from 60 to 80 m3/m2·h.
It may be advantageous, in a first reactor, to achieve a conversion of aromatic hydrocarbon of from 95 to 99.5% and, in a downstream reactor, a degree of conversion of >99.9%, in particular >99.99%, preferably >99.995%. In such a case, the ratio of the volumes of the catalyst beds of main reactor to downstream reactor is generally in the range from 20:1 to 3:1, in particular in the range from 15:1 to 5:1.
In a further embodiment, the main reactor can be operated in circulation mode. The circulation ratio (ratio of feed in kg/h to recycle stream in kg/h) is typically in the range from 1:5 to 1:100, preferably in the range from 1:10 to 1:50, preferentially in the range from 1:15 to 1:35, It is also possible in this case to remove the heat formed in the reaction partially or fully by passing the recycle stream through a heat exchanger.
In a further embodiment, the postreactor may also be integrated into the main reactor.
From case to case, it may also become necessary to regenerate the hydrogenation catalyst owing to declining activity. This is done by the methods which are customary for noble metal catalysts such as ruthenium catalysts and are known to those skilled in the art. These include, for example, the treatment of the catalyst with oxygen as described in BE 882 279, the treatment with diluted, halogen-free mineral acids as described in U.S. Pat. No. 4,072,628 or the treatment with hydrogen peroxide, for example in the form of aqueous solutions with a content of from 0.1 to 35% by weight, or the treatment with other oxidizing substances, preferably in the form of halogen-free solutions. Typically, the catalyst will be flushed with a solvent, for example water, after the reactivation and before the reuse.
The reaction product obtained in the process, i.e. the cycloaliphatic or the mixture of corresponding cycloaliphatics, can be purified further in a step c).
In the case that the reactant used is an aromatic hydrocarbon and the corresponding cycloaliphatic is obtained, the resulting reaction product can be subjected to a purifying distillation in order to remove any by-products formed, such as low boilers relative to the corresponding cycloaliphatic, for example n-hexane and n-pentane, or else high boilers. When, for example, benzene is used as the reactant, the cyclohexane obtained may comprise as impurities, for example, n-hexane and n-pentane, which can be removed as low boilers. Possible high boilers may include methylcyclohexane which can likewise be removed by distillation. In the purifying distillation, the pure cyclohexane can be obtained via a side draw in the column, while the low boiler components are drawn off at the top and high boiler components at the bottom. Alternatively, the purification of the product can also be effected in a column with a dividing wall, in which case the pure cyclohexane is drawn off at the level of the dividing wall.
When the reactant used is a mixture of aromatic hydrocarbons, the individual components of the cycloaliphatic mixture formed are separated by distillation and any further impurities are removed by distillation.
The heat of reaction arising in the course of the exothermic hydrogenation can, if appropriate, in the event of appropriate selection of the pressure level of the distillation, be utilized to operate the evaporator of the distillation column. To this end, the hot reaction effluent can be introduced directly into the column evaporator or, if appropriate, a secondary medium can be heated (for example generation of steam) and introduced into the column evaporator.
The partial steps of the process and also the overall process can be carried out continuously, semicontinuously or discontinuously.
With the aid of the process according to the invention, it is thus possible to obtain hydrogenated products which comprise very low residual contents, if any, of the starting materials to be hydrogenated.
The present invention further relates to a process for desulfurizing an aromatic hydrocarbon which comprises aromatic sulfur compounds, if appropriate in the presence of hydrogen, as described above in step a).
In hydrogenation processes in which the catalysts described above are used, deactivation is observed after a period of operation of the catalyst. Such a deactivated ruthenium catalyst can be brought back to the state of the original activity by flushing. The activity can be restored to >90%, preferably >95%, more preferably >98%, in particular >99%, most preferably >99.5%, of the original value. The deactivation is attributed to traces or residues of water adsorbed on the catalyst. This can surprisingly be reversed by flushing with inert gas. The regeneration method of the invention can thus also be referred to as drying of the catalyst or removal of water from this.
“Flushing” means that the catalyst is brought into contact with inert gas. Normally, the inert gas is then passed over the catalyst by means of suitable constructional measures known to those skilled in the art.
The flushing with inert gas is carried out at a temperature of from about 10 to 350° C., preferably from about 50 to 250° C., particularly preferably from about 70 to 180° C., most preferably from about 80 to 130° C.
The pressures applied during flushing are from 0.5 to 5 bar, preferably from 0.8 to 2 bar, in particular from 0.9 to 1.5 bar.
According to the invention, the treatment of the catalyst is preferably carried out using an inert gas. Preferred inert gases comprise nitrogen, carbon dioxide, helium, argon, neon and mixtures thereof. Nitrogen is most preferred.
In a particular embodiment of the invention, the inventive method of regeneration is carried out without removal of the catalyst in the same reactor in which the hydrogenation has taken place. The flushing of the catalyst according to the present invention is particularly advantageously carried out at temperatures and pressures in the reactor which correspond to or are similar to those in the hydrogenation reaction, resulting in only a very brief interruption of the reaction process.
According to the present invention, the flushing with inert gas is carried out at a volume flow of from 20 to 200 standard I/h, preferably at a volume flow of from 50 to 200 standard l/h per liter of catalyst.
The flushing with inert gas is preferably carried out for a time of from 10 to 50 hours, particularly preferably from 10 to 20 hours. For example, the calculated drying time of the catalyst bed of an industrial cyclohexane production plant having an assumed moisture content of 2 or 5% by weight is approximately 18 or 30 hours, respectively. The flushing according to the method of the invention can be carried out either in a downward direction (downflow mode) or in an upward direction (upflow mode).
The present invention further provides an integrated process for the hydrogenation of an aromatic hydrocarbon in the presence of a ruthenium catalyst having a catalyst regeneration step. In step a) of this process, the aromatic hydrocarbon or the mixture of aromatic hydrocarbons, each of which comprises aromatic sulfur compounds as an impurity, is desulfurized and hydrogenated in step b). Thereinafter the hydrogenating catalyst is regenerated by flushing with inert gas, as laid out above, until the original activity or part of the original activity is attained.
According to an embodiment of the invention the aromatic hydrocarbon is benzene. In a further embodiment the aromatic hydrocarbon is a mixture of benzene and Toluene or mixtures which comprise benzene and xylene or a xylene isomer mixture, or mixtures which comprise benzene, toluene and xylene or a xylene isomer mixture.
The method of the invention is also suitable for drying catalysts which have absorbed water during various procedures such as maintenance or storage.
The method of the invention is also suitable for drying catalysts which have absorbed water during various procedures such as maintenance or storage.
The invention will be illustrated hereinafter with reference to the examples adduced:
The experiments were performed in continuous tubular reactors with internal thermoelements (Ø 6 mm), trace heating (heating mats) and liquid metering.
The desulfurizing agent used was the catalyst R 3-12 from BASF Aktiengesellschaft in the form of 5×3 mm tablets—referred to hereinafter as catalyst A.
The desulfurizing agent was dried in accordance with the above description. To this end, the desulfurizing agent was heated to 200±10° C. in a nitrogen stream of 300±20 m3 (STP)/m3[CAT]·h at a heating rate not exceeding 50 K/h. As soon as the water had been removed, the desulfurizing agent was cooled to 120±5° C. at a cooling rate not exceeding 50 K/h. The drying procedure was effected in trickle mode (flow direction from the top downward).
In some cases, the desulfurizing agent was used in its reduced form. In this case, the desulfurizing agent was converted from its oxidized form to its reduced form with hydrogen in accordance with the description. To this end, the dried desulfurizing agent (in its oxidized form) was heated to 120±5° C. with a nitrogen stream of 300±20 m3 (STP)/m3[CAT]·h. 0.5±0.1% by volume of hydrogen was then metered to the abovementioned nitrogen stream until a temperature increase of from 15 to 20° C. occurred and remained constant. Subsequently, the hydrogen stream was increased to 1.0±0.1% by volume of hydrogen until a temperature increase of max. 30±5° C. occurred overall and the temperature again remained constant. The hydrogen stream was then increased to 2.0±0.2% by volume, in the course of which the temperature of the catalyst did not rise above 225° C. The hydrogen stream was then increased to 4.0±0.4% by volume and the temperature of the nitrogen was simultaneously increased to 200±10°, in the course of which the temperature of the catalyst did not rise above 225° C. A further increase in the hydrogen stream then to 6.0±0.6% by volume led to a rise in the temperature of the catalyst to 220±10° C., which was maintained. After one hour, the catalyst was then cooled to below 50° C. with a nitrogen stream of 300±2 m3 (STP)/m3[CAT]·h at a cooling rate not exceeding 50±5 K/h. The reaction procedure was effected in trickle mode (flow direction from the top downward).
The feedstock used was benzene with a purity of >99.95%.
The benzene used and the reaction effluents were analyzed by gas chromatography with reporting of GC area percentages (instrument: HP 5890-2 with autosampler; range; 4; column: 30 m DB1; film thickness: 1 μm; internal column diameter: 0.25 mm; sample volume: 5 μl; carrier gas: helium; flow rate, 100 ml/min; injector temperature: 200° C.; detector: FID; detector temperature: 250° C.; temperature program: 6 min at 40° C., 10° C./min to 200° C. for 8 min, total running time 30 min).
The total sulfur content in the benzene used and the reaction effluents were analyzed by Wickbold combustion by means of ion chromatography. To this end, from 4 to 6 g of the sample are mixed with acetone (Merck Suprasolv, item No. 1.0012.1000) in a ratio of 1:1 and then combusted in a hydrogen-oxygen gas flame in a Wickbold combustion apparatus. The combustion condensate is collected in an alkaline receiver which comprises 40 mmol of KOH (Merck Suprapure, item No. 1.050.020.500) (aqueous solution). The sulfate formed from the sulfur and collected in the receiver is determined by ion chromatography.
(ion chromatography system; modular system, from Metrohm; precolumn; DIONEX AG 12, 4 mm; separating column: DIONEX AS 12, 4 mm; eluent; 2.7 mM Na2CO3 (Merck Suprapure, item No. 1.063.950.500) and 0.28 mM NaHCO3 (Riedel de Haen, p.A., item No. 31437); flow rate: 1 ml/min; detection: conductivity after chemical suppression; suppressor: e.g. MSM, from Metrohm).
100 ml of catalyst A which had been dried by the drying procedure outlined above were charged in oxidic form into the above-described tubular reactor (Ø 25 mm×40 cm), the catalyst having been embedded into an inert bed of V4A rings above and below the actual catalyst bed. The height of the actual catalyst bed was approx. 22 cm. The experiment was carried out in liquid-phase mode at a pressure of 20 bar, 30 l (STP) of nitrogen per h having been metered into the liquid stream in cocurrent during the experiment.
The data compiled in Table 1 show clearly that the desulfurization of the benzene used can be carried out with catalyst A in oxidic form.
100 ml of catalyst A which had been dried in accordance with the drying procedure outlined above and reduced in accordance with the activation procedure outlined above were charged in reduced form into the above-described tubular reactor (Ø 25 mm×80 cm), the catalyst having been embedded into an inert bed of V4A rings above and below the actual catalyst bed. The height of the actual catalyst bed was approx. 22 cm. The experiment was carried out in liquid-phase mode at a pressure of 20 bar, a mixture of nitrogen and hydrogen having been metered into the liquid stream in cocurrent during the experiment.
The data compiled in Table 2 show clearly that the desulfurization of the benzene used can be carried out with catalyst A in reduced form.
100 ml of catalyst A which had been dried in accordance with the drying procedure outlined above and which had been reduced in accordance with the activation procedure outlined above were charged in reduced form into the above-described tubular reactor (Ø 25 mm×40 cm), the catalyst having been embedded into an inert bed of V4A rings above and below the actual catalyst bed. The height of the actual catalyst bed was approx. 22 cm. The experiment was carried out in liquid-phase mode at a pressure of 20 bar, 2 l (STP) of hydrogen per h having been metered into the liquid stream in cocurrent during the experiment.
The data compiled in Table 3 show clearly that the desulfurization of the benzene used can be carried out with catalyst A in reduced form even in extended operation. Moreover, the data show clearly that only very small amounts of benzene are reduced to cyclohexane.
After this extended experiment had ended, the spent catalyst was deinstalled and analyzed. To this end, the catalyst was oxidized slowly with a nitrogen/air mixture or with pure air at a temperature of approx. 25-30° C. The oxidized catalyst was deinstalled in ten separate fractions with approximately equal volumes, a sample was removed in each case and these were analyzed by elemental analysis. The result of the analysis is listed in Table 4. The samples are numbered in accordance with the flow direction (liquid-phase mode, fraction 1 at the bottom, fraction 10 at the top).
In accordance with expectation, the catalyst fraction at the reactor inlet (fraction 1) has the highest sulfur concentration, while the lowest content is present in the last fraction (fraction 10).
50 ml of catalyst A which had been dried in accordance with the drying procedure outlined above and which had been reduced in accordance with the activation procedure outlined above were charged in reduced form into the above-described tubular reactor (Ø 25 mm×40 cm), the catalyst having been embedded into an inert bed of V4A rings above and below the actual catalyst bed. The height of the actual catalyst bed was approx. 11 cm. The experiment was carried out in liquid-phase mode at a pressure of 3 bar, 2 l (STP) of hydrogen per h having been metered into the liquid stream in cocurrent during the experiment.
The data compiled in Table 5 show clearly that a desulfurization can be carried out at 3 bar, 80° C. and catalyst loading of >5 kgbenzene/Icatalyst·h.
A continuous tubular reactor (≡ 46 mm×3500 mm) was charged with 3700 ml of catalyst A, the catalyst having been embedded into an inert bed above and below the actual catalyst bed (800 ml and 500 ml respectively). The installed catalyst A was then dried and reduced in trickle mode in accordance with the procedure outlined in Table 6.
Subsequently, desulfurization was carried out at a pressure of from 3 to 32 bar in liquid-phase mode.
,The results of Table 7 show clearly that the content of aromatic sulfur compounds can be lowered below 70 ppb.
The experiment was performed in a continuous jacketed reactor, (Ø 12 mm×1050 mm) with three oil heating circuits distributed uniformly over the reactor length. The reactor was operated in continuous trickle mode with controlled liquid circulation (HPLC pump). The experimental plant was also equipped with a separator for separating gas and liquid phase with level control, offgas regulator, external heat exchanger and sampler. The hydrogen was metered under pressure control (in bar); the hydrogen used in excess was measured under quantitative control (in l (STP)/h); the benzene feedstock was metered via an HPLC pump. The product was discharged under level control via a valve. The temperature was measured with a thermoelement at the start (inlet) and at the end (outlet) of the reactor or of the catalyst bed. The benzene used had a total sulfur content of <0.1 mg/kg (detection by ion chromatography). The catalyst used was a meso-/macroporous Ru/Al2O3 catalyst with 0.47% by weight of Ru (catalyst B) or a mesoporous Ru/SiO2 catalyst with 0.32% by weight of Ru (catalyst C). These were prepared as detailed in the description. For example, catalyst C can be prepared as follows:
50 kg of the SiO2 support (D11-10 (BASF); 3 mm extrudates (No. 04/19668), water uptake of 0.95 ml/g, BET 135 m2/g) are initially charged in an impregnating drum and impregnated at 96-98% by weight water uptake. The aqueous impregnating solution comprises 0.176 kg of Ru as ruthenium acetate from Umicore, 4.34% by weight of Ru, batch 0255). The impregnated catalyst is dried without motion at an oven temperature of 145° C. down to a residual moisture content of approximately 1%. The reduction is effected with motion in hydrogen (approximately 75% H2 in N2, N2 being employed as the purge stream; 1.5 m3 (STP)/h of H2-0.5 m3 (STP)/h of N2) with a moving bed at 300° C. and a residence time of 90 minutes (1-2 h). The passivation is effected in dilute air (air in N2). The addition of air is controlled such that the temperature of the catalyst remains below 30-35° C. The finished catalyst C comprises 0.31-0.32% by weight of Ru.
This catalyst is described in detail below:
Tapped density of the shaped
The reduced catalyst C comprises at least partly crystalline ruthenium in the outermost zone (extrudate surface). In the support, ruthenium occurs in individual particles 1-10 nm (in places >5 nm): usually 1-5 nm. The size of the particles decreases from the outside inward.
Ruthenium particles are seen up to a depth of 30-50 micrometers below the extrudate surface. In this coating, ruthenium is present at least partly in crystalline form (SAD: selected area diffraction). The main portion of the ruthenium is thus in this coating (>90% within the first 50 μm).
A heatable 1.2 l pressure vessel (internal diameter 90 mm, vessel height: 200 mm, made of stainless steel) with 4-blade beam sparging stirrer, baffles and an internal riser for sampling or for charging and emptying the pressure vessel is charged with the particular amount (volume or mass) of the catalyst used in a “catalyst basket” (made of stainless steel).
The pressure vessel is sealed for pressure testing and charged with 50 bar of nitrogen. Afterward, the pressure vessel is decompressed, evacuated with a vacuum pump and isolated from the vacuum pump, and feedstock or the feedstock solution is sucked into the vessel via the riser.
To remove residual amounts of oxygen, the vessel is successively charged at room temperature twice with 10-15 bar each time of nitrogen and twice with 10-15 bar each time of hydrogen and decompressed.
The stirrer is switched on, a stirrer speed of 1000 rpm is established and the reaction solution is heated to reaction temperature. The target temperature is attained after 15 minutes at the latest. Hydrogen is injected up to the particular target pressure within 5 minutes. The hydrogen consumption is determined and the pressure is kept constant at the particular target pressure.
The riser is used at regular intervals to take preliminary samples (to flush the riser) and samples of the reaction mixture for monitoring the progress of the reaction.
After the appropriate reaction time, the heater is switched off, the pressure vessel is cooled to 25° C., the elevated pressure is released slowly and the reaction mixture is emptied via the riser with slightly elevated pressure. Afterward, the pressure vessel is evacuated with a vacuum pump and isolated from the vacuum pump, and new feedstock or the feedstock solution is sucked into the vessel via the riser.
This method enables the same catalyst to be used more than once. The hydrogen used had a purity of at least 99.9-99.99% by volume (based on dry gas). Secondary constituents are carbon monoxide (max. 10 ppm by volume), nitrogen (max. 100 ppm by volume), argon (max. 100 ppm by volume) and water (max. 400 ppm by volume).
The benzene used and the reaction effluents were analyzed by gas chromatography with reporting of GC area percentages (instrument: HP 5890-2 with autosampler; range: 4; column: 30 m DB1; film thickness: 1 μm; internal column diameter: 0.25 mm; sample volume: 5 μl; carrier gas: helium; flow rate: 100 ml/min; injector temperature: 200° C.; detector: FID; detector temperature: 250° C.; temperature program: 6 min at 40° C., 10° C./min to 200° C. for 8 min, total run time 30 min).
The total sulfur content in the benzene used and the reaction effluents were analyzed by Wickbold combustion by means of ion chromatography. To this end, from 4 to 6 g of the sample are mixed with acetone (Merck Suprasolv, item No. 1.0012.1000) in a ratio of 1:1 and then combusted in a hydrogen-oxygen gas flame in a Wickbold combustion apparatus. The combustion condensate is collected in an alkaline receiver which comprises 40 mmol of KOH (Merck Suprapure, item No. 1.050.020.500) (in water). The sulfate formed from the sulfur and collected in the receiver is determined by ion chromatography.
(Ion chromatography system: modular system, from Metrohm; precolumn: DIONEX AG 12, 4 mm; separating column: DIONEX AS 12, 4 mm; eluent: 2.7 mM Na2CO3 (Merck Suprapure, item No. 1.063.950.500) and 0.28 mM NaHCO3 (Riedel de Haen, p.A., item No. 31437); flow rate: 1 ml/min; detection: conductivity after chemical suppression; suppressor: e.g. MSM, from Metrohm).
104 ml (63.9 g) of catalyst B were used for continuous hydrogenation at a hydrogen pressure of 32 bar, at an offgas rate of 1-5 l (STP)/h, a reactor input temperature of 88-100° C. and a feed/circulation ratio of 1:30. The results are compiled in Table 8.
104 ml (63.9 g) of catalyst B were used for continuous hydrogenation at a hydrogen pressure of 19 bar, at an offgas rate of 1-5 l (STP)/h, a reactor input temperature of 88-100° C. and a feed/circulation ratio of 1:30. The results are compiled in Table 9.
104 ml (45.0 g) of catalyst C were used for continuous hydrogenation at a hydrogen pressure of 32 bar, at an offgas rate of 1-5 l (STP)/h, a reactor input temperature of 88-100° C. and a feed/circulation ratio of 1:30. The results are compiled in Table 10.
a) [GC area %];
b) [GC area-ppm]
a) [GC area %];
b) [GC area-ppm]
a) [GC area %];
b) [GC area-ppm]
The data compiled in Tables 8 to 10 show clearly that cyclohexane can be obtained with an excellent selectivity.
The hydrogenation plant consists of a storage tank for the desulfurized benzene, a reservoir vessel, a metering pump for benzene, a main reactor (Ø 45×2000 mm) with separator for separating gas and liquid, and regulator for level control, liquid circulation with pump and a heat exchanger for removing the heat of reaction formed, a postreactor (Ø 22 mm×1500 mm) with a separator for separating gas and liquid, and regulator for level control, and also a storage tank for the hydrogenation effluent. The main reactor and postreactor were each equipped with an internal thermoelement (Ø 6 mm in the main reactor, Ø 3 mm in the postreactor). Both reactors were operated in trickle mode. Liquid and gas were metered in in cocurrent.
The main reactor was charged with 2700 ml (1870 g), the postreactor with 340 ml (229 g), of catalyst B. For insulation and trace heating, the main reactor was equipped with electrical heating mats. The postreactor was manufactured for an adiabatic operating mode and was provided with appropriate insulation. Above and below the catalyst, an inert bed was introduced (wire mesh rings of stainless steel).
The feedstock used was desulfurized benzene which had been prepared analogously to Example a4 or a5, and had a total sulfur content of <0.1 mg/kg.
The benzene and the cyclohexane were analyzed by gas chromatography with reporting of GC area % or GC area-ppm; the analyses were carried out without internal standard (instrument: HP 5890-2 with autosampler; range: 4; column: 30 m DB1; film thickness: 1 μm; internal column diameter: 0.25 mm; sample volume: 5 μl; carrier gas: helium; flow rate: 100 ml/min; injector temperature: 200° C.; detector: FID; detector temperature: 250° C.; temperature program: 6 min at 40° C., 10° C./min to 200° C. for 8 min, total run time 30 min).
The total sulfur content in the benzene used and the reaction effluents were analyzed by Wickbold combustion by means of ion chromatography. To this end, from 4 to 6 g of the sample are mixed with acetone (Merck Suprasolv, item No. 1.0012.1000) in a ratio of 1:1 and then combusted in a hydrogen-oxygen gas flame in a Wickbold combustion apparatus. The combustion condensate is collected in an alkaline receiver which comprises 40 mmol of KOH (Merck Suprapure, item No. 1.050.020.500) (aqueous solution). The sulfate formed from the sulfur and collected in the receiver is determined by ion chromatography.
(Ion chromatography system: modular system, from Metrohm; precolumn. DIONEX AG 12, 4 mm; separating column: DIONEX AS 12, 4 mm; eluent: 2.7 mM Na2CO3 (Merck Suprapure, item No. 1.063.950.500) and 0.28 mM NaHCO3 (Riedel de Haen, p.A., item No. 31437); flow rate: 1 ml/min; detection: conductivity after chemical suppression; suppressor: e.g. MSM, from Metrohm).
In some cases, the sulfur content was determined by gas chromatography (detection limits in each case 50 ppb for COS and thiophene) (separating column: CP SIL88 (100% cyanopropylpolysiloxane), length: 50 m; film thickness: 0.2 μm; internal diameter: 0.25 mm; carrier gas: helium; initial pressure: 1.5 bar; split: on column (ml/min); septum purge: 5 ml/min; oven temperature: 60° C.; preheating time: 10 min; rate 1: 5° C./min; oven temperature 1: 200° C.; continued heating time 1: 10 min; rate 2: —; oven temperature 2: —; continued heating time 2: —; injector temperature: on column (° C.); detector temperature: 220° C.; injector: HP autosampler; injection volume: 1.0 μl; detector type: PFPD (flame photometer); GC method: % by weight method with external calibration; special features: ON-column injection and special flame photometer detector).
At the start, the plant was operated at 20 bar; the plant pressure was increased to 32 bar after 860 operating hours. In the downstream reactor, hydrogenation was effected up to full conversion; in the reaction effluent, virtually no benzene was detectable any longer.
The present results show that the process according to the invention enables benzene to be converted fully and cyclohexane to be obtained in high purities.
The experiment was carried out under the same conditions and in the same plant as described in Example b4). However, the main reactor was charged with 2700 ml (1218 g) of catalyst C and the postreactor with 340 ml (153 g) of catalyst C. In addition, the plant was operated at 32 bar from the start onward.
The results obtained here too show that virtually no benzene is detectable any longer in the reaction effluent.
Additionally metered into the feed after a run time of 5347 h were 4.3% by weight of toluene. The corresponding amounts of methylcyclohexane were found in the reaction effluent but no toluene.
750 ml of a 5% solution of benzene in cyclohexane were hydrogenated with 9.0 g (approx. 22 ml) of catalyst C at a temperature of 100° C. and a pressure of 20 bar with hydrogen. The catalyst was used repeatedly in five successive experiments. Samples were taken after reaction times of 10, 20, 30, 40, 60, 90, 120 and 180 minutes.
Table 13 lists the decrease in the benzene content over time. The mean values of the results of the five experiments and the maximum positive and negative deviation from the mean for the particular samples are evaluated. The benzene content was determined by means of GC analysis in GC area %,
750 ml of a 5% solution of benzene in cyclohexane were hydrogenated with 9.0 g (approx. 22 ml) of catalyst C at a temperature of 100° C. and a pressure of 32 bar with hydrogen. The catalyst was used repeatedly in five successive experiments. Samples were taken after reaction times of 10, 20, 30, 40, 60, 90, 120 and 180 minutes.
Table 14 lists the decrease in the benzene content over time. The mean values of the results of the five experiments and the maximum positive and negative deviation from the mean for the particular samples are evaluated. The benzene content was determined by means of GC analysis in GC area %.
A mesoporous/macroporous aluminum oxide support in the form of 3-5 mm sphere having a total volume of 0.44 cm3/g, with 0.09 cm3/g (20% of the total pore volume) being formed by pores having a diameter in the range from 50 nm to 10 000 nm and 0.35 cm3/g (80% of the total pore volume) being formed by pores having a diameter in the range from 2 nm to 50 nm, a mean pore diameter in the region of 11 nm and a surface area of 286 m2/g was impregnated with an aqueous ruthenium(III) nitrate solution. The volume of solution taken up during impregnation corresponded approximately to the pore volume of the support used. The support impregnated with the ruthenium(III) nitrate solution was subsequently dried at 120° C. and activated (reduced) in a stream of hydrogen at 200° C. The catalyst produced in this way comprised 0.5% by weight of ruthenium, based on the weight of the catalyst. The ruthenium surface area was 0.72 m2/g, and the ratio of ruthenium surface area to support surface area was 0.0027.
The affinity of the catalyst for water was determined by means of measurements of the sorption of water vapor on the catalyst produced as described above (0.5% Ru/γ-Al2O3).
It was found that the catalyst sorbs an amount of water of 5% even at relatively low vapor pressures of 30%. If only traces of water are present in the reactor or in the starting materials, this water can be sorbed on the catalyst.
In a plant for the preparation of cyclohexane using a ruthenium/aluminum oxide catalyst comprising 0.5% of Ru on a γ-Al2O3 support, a steady decrease in the catalyst activity and an increasing benzene content in the product stream are observed. Further monitoring of the reaction during a catalyst operating life test shows that the residual benzene content downstream of the main reactor in the hydrogenation of benzene increases from a few hundred ppm to some thousands of ppm over a period of operation of about 3400 hours. A calculation indicates that introduction of 16 620 kg/h of benzene having a water content of from 30 to 50 ppm introduces 0.8 kg of water per hour into the plant. In addition to this, there are a further 3.5 kg/h of water originating from the hydrogen gas.
When the plant was shut down after 3394 hours of operation, the plant ran with a residual benzene content of 0.2% at a WHSV of 0.6 gbenzene/mlcat·h. During shutdown, the plant was flushed with pressurized nitrogen at a temperature of 70-100° C. and then depressurized. After start-up, the plant gave a residual benzene content of from 0.01% to 0.04% at a WHSV of 0.6 gbenzene/mlcat·h.
This observed effect of drying of the catalyst was verified again after 7288 hours of operation. At a WHSV of 0.9 gbenzene/mlcat·h, the residual benzene content at the end of the plant was 0.2% and even rose to 0.56%. After shutdown of the plant, the catalyst was dried by means of 100 standard l/h of nitrogen at 110° C. for a period of 34 hours. After start-up of the plant at a WHSV of 0.6 gbenzene/mlcat·h, the residual benzene content was from 0.03% to 0.07%, which can be attributed to a significant increase in the catalyst activity as a result of drying.
In both cases, drying of the catalyst led to a significantly higher catalyst activity which is close to or equal to the original catalyst activity.
To simulate the influence of water on the hydrogenation of benzene using a ruthenium catalyst, series of autoclave experiments before and after saturation of the catalyst with water and after drying of the catalyst were carried out. A 5% strength solution of benzene in cyclohexane together with the ruthenium catalyst was placed in the pressure vessel, the mixture was heated to the reaction temperature of 100° C. and the course of the reaction at a hydrogen pressure of 32 bar was followed by regular sampling. The samples were subsequently analyzed by gas chromatography.
23 hydrogenation experiments were carried out, and the catalyst was subsequently placed in water. 13 further hydrogenation experiments were then carried out. The catalyst displayed a significantly lower but virtually constant activity. After drying of the catalyst in a stream of nitrogen at 100° C. in a reaction tube, 5 further experiments were carried out; the catalyst displayed a hydrogenation activity similar to that before saturation with water.
The experiments demonstrate that the activity of the ruthenium/aluminum oxide catalyst used decreases significantly after contact with water, but the catalyst can be reactivated again by drying in a stream of nitrogen and the initial activity can be virtually fully restored.
| Number | Date | Country | Kind |
|---|---|---|---|
| 10 2005 062 354.9 | Dec 2005 | DE | national |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/EP2006/070186 | 12/22/2006 | WO | 00 | 6/23/2008 |
