KR100403387B1 - Process for preparing tertiary alkyl ethers from olefinic hydrocarbon feedstock - Google Patents

Abstract

The present invention relates to a process for the preparation of tertiary alkyl ethers.

The process selectively contacts at least a portion of the polyunsaturated hydrocarbons by contacting feedstock containing saturated or monounsaturated or polyunsaturated hydrocarbons containing C ₄ to C ₇ with excess hydrogen to the polyunsaturated hydrocarbons in the presence of a catalyst. Hydrogenating to produce a modified hydrocarbon feedstock wherein the unsaturated hydrocarbon consists predominantly of monounsaturated compounds; Supplying the modified hydrocarbon feedstock with at least a portion of the unreacted hydrogen to a catalytic distillation reactor system comprising at least one distillation column; Reacting the C _4-7 isoolefin in the feedstock in the presence of a catalyst to produce a tertiary alkyl ether; Distilling the reaction mixture in a distillation column; Recovering all non-reacted hydrocarbons with alkyl ether, and bottom product of distillation; Evacuating the ceiling product of the distillation, which contains primarily an azeotropic mixture of saturated C ₄ hydrocarbons and alkanols, and controlling the amount of gaseous distillation containing hydrogen in the column to substantially maintain the pressure in the distillation column. Consists of steps.

The selective hydrogenation operation of the present invention reduces the amount of diene contained in the feed to 0.5% or less, and also causes isomerization of the olefins, thereby increasing the amount of reactive olefins.

Description

Process for preparing tertiary alkyl ethers from olefinic hydrocarbon feedstock

The present invention relates in particular to a process for producing tertiary alkyl ether products for use as motor fuel components. These products contain, for example, methyl t-butyl ether, ethyl t-butyl ether, t-amyl methyl ether or t-amyl ethyl ether or mixtures thereof, and possibly high molecular weight tertiary alkyl ethers. According to the process of the invention, the isoolefins of the olefinic hydrocarbon feedstock, in particular the C ₄ -C ₇ isoolefins, are reacted with a suitable alkanol to produce the corresponding ethers. This ether is separated with the bottom product of the distillation reaction system and further processed as necessary to produce the motor fuel component. Unreacted alkanol is separated together with the ceiling product of distillation.

Tertiary alkyl ethers are added to the gasoline to improve their anti-detonation properties and to reduce the concentration of harmful components in the exhaust gas. Oxygen-containing ether groups of the compounds have been found to advantageously improve the combustion process of automotive engines. Examples of suitable alkyl tertiary alkyl ethers are methyl t-butyl ether (MTBE), ethyl t-butyl ether (ETBE), t-amyl methyl ether (TAME), t-amyl ethyl ether (TAEE) and t-hexyl methyl ether (THME). These ethers are prepared by etherification of isoolefins with monoaliphatic alcohols (alkanols). The reaction can be carried out in a fixed bed reactor, fluidized bed reactor, tubular reactor or in a catalytic distillation column.

The etherification reaction is an exothermic equilibrium reaction and the maximum conversion is determined by the thermodynamic equilibrium of the reaction system. To use TAME as an example, the reaction and separation in a reaction distillation column can be performed to achieve about 90% conversion, while in a fixed bed reactor only 65% to 70% conversion can be achieved.

Ion exchange resins are the most common etherification catalysts. Generally, the resins used include cation exchange resins (sulfonated polystyrene crosslinked with divinylbenzene) based on sulfonated polystyrene / divinylbenzene having a particle size of 0.1 to 1 mm.

Two TAME processes have been used commercially for some time. The first method comprises a fixed bed reactor, a column for product separation by distillation, and a methanol separation device. In another method, product distillation is replaced by catalytic distillation apparatus, which significantly improves TAME conversion as mentioned above.

A third completely novel etherification process is described in the applicant's international patent application WO 93/19031. This novel process involves a catalytic distillation apparatus modified by transferring the catalyst normally positioned in the distillation column into a separate external reactor fed from the product separation distillation apparatus. The side reactor product is recycled to the same product separation distillation apparatus. According to one embodiment of the process described in the applicant's international patent application WO 93/19032, the product distillation of the catalytic distillation reactor system is characterized in that most, preferably substantially all of the alkanols which are separated together with the distillate are inert C ₄ in combination with a hydrocarbon to produce an azeotrope with it. The product is recovered from the bottom of the column and predominantly comprises a mixture of TAME and high molecular weight ethers.

The method described in the applicant's international patent application described above can also be used to prepare lower alkyl ethers such as methyl t-butyl ether (MTBE) and ethyl t-butyl ether (ETBE) and mixed ether products containing these ethers. Can be. Suitable feedstocks for the method for producing a tertiary alkyl ether is a considerable amount, typically not less than 5% by weight, typically from C _4-7 comprising a reactive C _4-7 iso-olefin of about 7 to 30% by weight Hydrogenated catalyst cracking (FCC) gasoline containing hydrocarbons. These reactive isoolefins are isobutene, 2-methyl-1-butene, 2-methyl-2-butene, 2-methyl-1-pentene, 2-methyl-2-pentene, 2,3-dimethyl-1-butene, 2,3-dimethyl-2-butene, 2-ethyl-1-butene, 2-methyl-2-hexene, 2,3-dimethyl-1-pentene, 2,3-dimethyl-2-pentene, 2,4- Dimethyl-1-pentene, 2-ethyl-1-pentene and 2-ethyl-2-pentene.

However, FCC gasoline also contains about 1 to 3 weight percent diene, such as butadiene, isoprene, pentadiene and cyclopentadiene. These dienes are polymerized in the presence of ion exchange resins of the type used as catalysts in the etherification process. As a result of the polymerization of the diene, rubber is produced that blocks the pores of the ion exchange catalyst and reduces the activity of the catalyst. In automotive engines, these rubbers are also not beneficial because they clog the engine valves. Another problem associated with diene compounds is their color, which is not acceptable for motor gasoline.

Other hydrocarbon feedstocks suitable for the etherification process are produced by pyrolysis C ₅ gasoline, Thermofor Catalytic Cracking (TCC) gasoline, residual catalytic cracking (RCC) gasoline and coker gasoline. These feedstocks also contain harmful diene compounds and are therefore limited by the same problems as FCC gasoline.

Summary of the Invention

It is an object of the present invention to solve the problems associated with the prior art by providing a novel process for producing tertiary alkyl ethers from olefinic hydrocarbon feedstocks containing more than about 0.5% diene.

The present invention is mainly based on the novel etherification process described in international patent applications WO 93/19031 and WO 93/19032. In particular, the hydrocarbon feedstock and the at least one alkanol are fed into a catalytic distillation reactor system comprising one or more distillation columns. In a catalytic distillation reactor system, the components of the feed, ie alkanols and reactive isoolefins, react with each other to produce a product containing tertiary alkyl ethers. The reaction mixture is distilled continuously in the distillation column of the system. The bottom product containing mainly the resulting alkyl ethers and virtually all unreacted unsaturated hydrocarbons is recovered from the distiller, the ceiling products being azeotropic mixtures produced by unreactive (inactive) feed hydrocarbons, in particular unreactive saturated C ₄ hydrocarbons, and It contains mainly alkanols which are not consumed by the etherification reaction.

According to the present invention, the hydrocarbon feedstock is modified by a selective hydrogenation reaction before being fed into the catalytic distillation reactor system. As a result of the selective hydrogenation, a significant amount of harmful diene contained in the olefinic hydrocarbon feedstock is removed and a modified hydrocarbon feedstock containing less than about 0.5% by weight diene is obtained. The selective hydrogenation reaction according to the invention is carried out in the presence of a suitable selective hydrogenation catalyst using excess hydrogen. Importantly, after hydrogenation, excess hydrogen, ie unreacted hydrogen, is not separated from the product stream, but rather, with the modified hydrocarbon feedstock, is led to a reactive distillation reactor system and used for the control and control of distillation.

In particular, the method according to the invention is mainly characterized by the following steps:

Optionally in the presence of a catalyst, a saturated hydrocarbon having 4 to 7 carbon atoms, and a feedstock containing monounsaturated and polyunsaturated hydrocarbons, is contacted with excess hydrogen to the polyunsaturated hydrocarbons to selectively select some or all of the polyunsaturated hydrocarbons. Hydrogenation to produce a modified hydrocarbon feedstock in which the unsaturated hydrocarbons consist primarily of monounsaturated compounds,

Supplying some or all of the unreacted hydrogen and the modified hydrocarbon feedstock to a catalytic distillation reactor system comprising at least one distillation column,

Reacting the C _4-7 isoolefin in the feedstock with an alkanol in the presence of a catalyst to produce a tertiary alkyl ether,

Distilling the reaction mixture in a distillation column,

Recovering virtually all unreacted hydrocarbons with alkyl ether and bottom product of distillation,

Draining the ceiling product of distillation, which mainly contains an azeotropic mixture of C ₄ hydrocarbons and alkanols, and

Maintaining a substantially constant pressure in the distillation column by adjusting the amount of gaseous distillate containing hydrogen in the distillation column.

Within the scope of the present application, the expression “catalytic distillation reactor system” refers to an apparatus in which the ether product reaction and separation of the product are performed at least partially simultaneously. Such apparatus includes a conventional reactive distillation column or a distillation column in combination with at least one side reactor. Reference is made to implementations described in detail in international patent applications WO 93/19031 and WO 93/19032.

The term "alkanol" includes lower alkyl alcohols capable of forming an azeotrope with saturated or unsaturated hydrocarbons, in particular C ₃ to C ₇ hydrocarbons, of the hydrocarbon feedstock. As specific examples of alkanols, mention may be made of methanol, ethanol, n-propanol, i-propanol, n-butanol, i-butanol and t-butanol. Methanol and ethanol are particularly preferred.

Preferred "mono unsaturated" hydrocarbons (ie hydrocarbons containing only one double bond) are isobutene, 2-methyl-1-butene, 2-methyl-2-butene, 2-methyl-1 -Pentene, 2-methyl-2-pentene, 2,3-dimethyl-1-butene, 2,3-dimethyl-2-butene, 2-ethyl-1-butene, 2-methyl-2-hexene, 2,3 -Dimethyl-1-pentene, 2,3-dimethyl-2-pentene, 2,4-dimethyl-1-pentene, 2-ethyl-1-pentene, 2-ethyl-2-pentene and 1-methyl-cyclopentene Include.

Representative examples of “polyunsaturated” hydrocarbons described herein are mainly butadiene, pentadiene (eg isoprene) and cyclopentadiene, and dienes having six or more carbon atoms. Selective hydrogenation of similar higher dienes containing isoprene and methyl side chains is particularly important because they can produce isoolefins that can be etherified. In addition, acetylene-based compounds such as butyne, in particular 2-butyne, fentin and hexine, are also present in the olefinic hydrocarbon feedstock. Such compounds will produce monounsaturated hydrocarbons in the hydrogenation reaction.

The term "olefinic hydrocarbon feedstock" includes all hydrocarbon feedstocks that contain a mixture of isoolefins that can be etherified to produce tertiary alkyl ethers. In particular, feedstocks such as FCC gasoline, FCC light gasoline, pyrolysis C ₅ gasoline, TCC gasoline, RCC and coker gasoline are preferred. The feed also includes a mixture of two or more olefinic hydrocarbon feedstocks, such as a mixture of FCC light gasoline and pyrolysis C ₅ cuts. Of course, the proportion of various C ₄ to C ₇ isoolefins will have a greater impact on the composition of the ether product.

Of these feedstocks, FCC, RCC and TCC are preferred because such hydrocarbon cuts can be used on their own, possibly after removal of high molecular weight cuts (C ₈₊ ). Pyrolysis gasoline has to be removed of low molecular weight cuts and C ₆₊ cuts before being fed into the hydrogenation unit. In particular, it is important to remove all aromatic compounds. Up to about 10% of the C ₆₊ cuts are included in the resulting hydrocarbon mixture called pyrolysis C ₅ gasoline to ensure that virtually all reactive C ₅ of the pyrolysis gasoline is present in the olefinic feedstock. This feedstock will also contain reactive aliphatic C ₆₊ hydrocarbons. Pyrolysis gasoline is particularly rich in isoprene (up to 10% by weight), and the selective hydrocarbons according to the invention, in combination with any of the above mentioned cracking gasoline cuts, will greatly improve the value of the cuts as a feedstock for etherification. will be.

Selective hydrogenation of pure C ₅ cut dienes is known in the art (see, eg, US Pat. Nos. 4,724,274 and 5,352,846). These cuts are chemically more uniform and the molecular weights of the hydrocarbons contained therein have the same size. In contrast, according to the present invention, the olefinic hydrocarbon feedstock to be hydrogenated contains a complex mixture of a wide range of C ₄ to C ₇ hydrocarbons. These mixtures comprise 10 different aliphatic dienes which may be cyclic as well as straight and branched chains. An example of the species that may be mentioned last is olefinic naphthenes. Thus, it is surprising that the conventional hydrogenation catalyst used in the present invention can achieve selective hydrogenation of all the different dienes present in the mixture of C ₄ to C ₇ hydrocarbons, thereby reducing the diene content to less than 0.1% by weight. The greater the reactivity of the diene, the easier the hydrogenation becomes. In the following, the hydrogenation reaction carried out in accordance with the invention called "selective hydrogenation" is a polyunsaturated compound in which such hydrogenation is contained in the hydrocarbon feedstock, even though a small amount of the monounsaturated component originally contained in the feed may be saturated. Influences to convert them into monounsaturated compounds.

The hydrogenation process of the present invention not only achieves selective hydrogenation of the feed, but also isomerizes the olefin compounds. In the etherification reaction the reactive component is a tertiary olefin. Among tertiary olefins, alpha-olefins with double bonds closer to the ends of the molecule are more reactive than beta-olefins. As far as the etherification of the olefinic feedstock containing the mixture of reactive olefins is concerned, it is very important to maximize the amount of alpha-olefins produced. Surprisingly, the selective hydrogenation process of the present invention also isomerizes some non-reactive components such as 3-methyl-1-butene and isoprene with reactive components such as 2-methyl-1-butene and 2-methyl-2-butene Has been proven. Isomerization significantly increases the profitability of etherification.

Hydrogenation produces a modified hydrocarbon feedstock, wherein at least about 80%, preferably at least about 90%, in particular at least 99.9% of its unsaturated hydrocarbons consist of monounsaturated compounds. Of the polyunsaturated hydrocarbons, at least about 80%, preferably at least 90% are hydrogenated during the selective hydrogenation operation.

1 is a schematic diagram of a method according to the invention. As such, in the etherification process of the present invention, the olefinic hydrocarbon feedstock, for example FCC gasoline, optionally removes the high molecular weight fraction in the FCC gasoline splitter 1 to produce FCC light gasoline, followed by a selective hydrogenation device. Passed to (2), if present, produces a modified hydrocarbon feedstock containing only a small amount of diene compound. A modified feedstock containing C _4-7 hydrocarbons is mixed with a suitable alkanol and fed into an etherification apparatus 3 comprising a catalytic distillation reactor system. In the etherification unit, the C _4-7 isoolefins of the feedstock react with the alkanol in the presence of a cation exchange resin to produce a tertiary alkyl ether. The alkyl ether is recovered from the distillation reactor system with the bottom product and, if desired, further processed to produce etherified gasoline product. Most of the unreacted hydrocarbons (except C ₄ ) are also recovered with the bottom product of the distillation.

According to a preferred embodiment, the selective hydrogenation device consists of two sequentially arranged reactors. As far as the basic concept of the present invention is concerned, satisfactory results are obtained with one hydrogenation reactor. However, two or more reactors are preferred in order to lengthen the operating cycle sufficiently. As such, when the activity of the catalyst in the first selective hydrogenation reactor is reduced, a second reactor can be used and the production can continue. The cost of the reaction vessel along with the necessary equipment is less than the cost of production loss. Moreover, the use of two reactors reduces the catalyst requirement, since only the reactor in which it is operated needs to be filled with catalyst. If one large reactor is used, most of the catalyst will remain at the beginning of operation and will incur unnecessary costs.

FIG. 2 shows a preferred configuration of a selective hydrogenation apparatus comprising a first hydrogenation reactor 4 and a second hydrogenation reactor 5. The hydrocarbon feedstock and the hydrogen gas are combined before entering the first reactor and are led to a hydrogenation apparatus under pressure. The feed conduit may be equipped with a static mixer to facilitate the mixing of gaseous hydrogen into the liquefied hydrocarbon feed. Dissolution of hydrogen onto the hydrocarbon and hydrogenation of the diene are achieved by allowing the hydrogen to be evenly dispersed in the hydrocarbon feed in the form of fine bubbles.

The hydrogen used for the hydrogenation may consist of hydrogen gas of any suitable purity (typically 10 to 100% by weight). The hydrogen gas contains 0 to 90% by weight, preferably 10 to 70% by weight, of other volatile components (eg hydrocarbons), which remain inert during the selective hydrogenation and etherification reactions. Hydrocarbons of this kind can be used with excess hydrogen to control the pressure of the column. Example 1 shows the use of hydrogen gas containing up to about 50-60% by weight of methane and other low molecular weight paraffinic hydrocarbons.

Selective hydrogenation is operated at a pressure of 10 to 120 bar, preferably 15 to 115 bar, in particular about 20 to 40 bar to keep the hydrogen in the liquid state. It is preferred that a substantial amount, suitably at least about 80%, in particular at least 90%, of hydrogen be dissolved in the hydrocarbon feedstock during hydrogenation.

Hydrogen is supplied in a molar ratio of about 1.001 to 10 times, preferably 1.05 to 2 times, the amount of diene present in the feed. When calculating the molar amount of the diene, any acetylene hydrocarbon should be included. In order to limit hydrocarbon loss during pressure control in the column (see below), some excess hydrogen is preferred. Accordingly, the preferred range of hydrogen described above, which corresponds to an excess of 1 to 1000 mol% relative to the diene of the hydrocarbon feedstock, is suitable. Especially preferred is an excess of 5 to 200 mol%, for example 10 to 40 mol%.

After hydrogenation, excess hydrogen is fed into the etherification reactor apparatus. According to one preferred embodiment, all unreacted hydrogen, ie essentially all excess hydrogen used for hydrogenation, is led to an etherification unit together with the selectively hydrogenated hydrocarbon feedstock. By introducing all excess hydrogen into the etherification unit, no separate hydrogen removal facility is required. However, if desired, it is also possible to remove some, for example 10 to 90% of unreacted hydrogen prior to etherification.

The catalyst for the selective hydrogenation reaction according to the invention preferably comprises a supported precious metal (group VIII) hydrogenation catalyst. Nickel-based catalysts may also be used. Catalyst metals include Pd, Pt, Re, Rh, Ag and Au, and mixtures thereof. The catalyst support is typically composed of an inorganic metal oxide such as alumina or silica. Typical examples of particularly suitable catalysts are commercially available hetero palladium based hydrogenation catalysts. The amount of palladium is generally 0.2 to 5% by weight, based on the total weight of the catalyst. The support consists of alumina.

If the action of the catalyst needs to be controlled, small amounts of sulfur can be added during the hydrogenation reaction. In this case, the sulfur compound is used in an amount of 1 to 100 ppm when calculated based on the hydrocarbon charge.

The hydrogenation reaction temperature is in the range of about 20 to 300 ° C. The temperature depends on the pressure, catalyst and hydrogen to hydrocarbon ratio. According to a preferred embodiment, the temperature increases from about 40 ° C. to about 130 ° C. during the operation, to compensate for the decrease in activity of the catalyst. As such, when using a supported palladium catalyst and working with 10% more excess hydrogen at a pressure of about 20 to 40 bar, the initial temperature can be set at 70 to 80 ° C. and the final temperature is 120 to May be set at 130 ° C.

According to a preferred embodiment of the invention, at least 5 mole% (eg 10 to 15 mole%) excess hydrogen is fed to the hydrogenation process, and the pressure of the hydrogenation process is maintained at 15 bar or more. In this case, the distillation pressure is kept below 10 bar.

3 shows a simplified schematic diagram of the TAME process described in the Examples.

The selectively hydrogenated hydrocarbon feedstock and the methanol required for etherification are fed under pressure (pumps 29 and 30), mixed with each other, and the mixture is heated or cooled to obtain the appropriate feed temperature, and then the mixture is pre-reactor (6). , 7 and 8). It should be noted that according to the invention, it is possible to feed alkanol into a hydrocarbon feedstock or into a substream entering the side reactor or to feed alkanol in a combination of these two alternatives. The preliminary reactor consists of three reactors filled with ion exchange resins. Such reactors may be fixed bed reactors, fluidized bed reactors or tubular reactors. The reactors can be arranged in line or in parallel as shown in the figure. If there are more than two preliminary reactors, they may also be arranged in line / parallel. Due to the reaction, the temperature of the preliminary reactor rises in the range of about 5 ° C. to 15 ° C., depending on the reactor insulation efficiency. Heat exchangers 16 to 18 are placed in front of the preliminary reactors 6 to 8 to heat or cool the fluid stream to the desired temperature.

From the preliminary reactor, the mixture is led via a further heat exchanger 19 to a distillation column 9, which is also referred to as the main fractionator. The distillation column may be a packed column or a column equipped with a valve, sieve or bubble-capture tray. At the bottom of the column, a reboiler 20 is installed. The ceiling gas of the column is discharged to the reflux drum 11 via the condenser 10, and the ceiling gas is removed by the pump 12 from the reflux drum. Some of the ceiling gases are further processed, for example provided to the MTBE process, and some of them are returned to the column. TAME and high molecular weight ethers are recovered with the bottom product using pump 28 and, if desired, cooled (27) before being led to the reservoir. The bottom product contains unreacted C ₅₊ hydrocarbons in addition to ether.

The reflux ratio of the column is preferably about 1/2 to 200. Even larger ratios can be used in experimental plant equipment. According to the invention, the reflux ratio is adjusted such that the amount of distillate removed from the process corresponds substantially to the amount of C ₄ hydrocarbons and other inert feed hydrocarbons which results in an azeotropic mixture with the alkanol used for etherification. At the top of the column is provided a pressure regulating means 21, which is operatively connected to a flow regulating means 22 for regulating the amount of gaseous distillate removed from the reflux drum 11.

Following the distillation column 9 was arranged a side reactor system consisting of a series of two reactors 13, 14. This reactor can be replaced with one large reactor, if desired. The substream may flow by the use of a pump or as a forced circulation by thermosiphon. Depending on the circulation scheme, the reactor may be a fixed bed, fluidized bed or tubular reactor, as described above. The side reactor is fed with a liquid stream taken from the column. The pressure of the liquid stream can be increased by the pump 23. The substream is preferably taken from a tray located below the tray having a methanol K-value of less than one. Additional methanol may be fed to the subreactor feed 15 in front of the subreactor if desired. The reactor feed may be cooled to the reaction temperature before the subreactor by heat exchanger 24. The temperature rises several degrees in the subreactor. The temperature of the liquid streams from the first and second reactors can be controlled by heat exchangers 25 and 26, respectively. From the subreactor system 13, 14, the liquid flows back to the column 9. The liquid is then returned to the plate with a methanol K-value greater than one.

The side reactor effluent typically enters the column to a position below the feed position derived from the preliminary reactors 6-8. The purpose of this arrangement is to operate the column 9 in such a way that the methanol in the ceiling product is bound to the C ₄ hydrocarbon in the form of an azeotrope.

Other alkanol ethers such as MTBE, ETBE and TAEE can be prepared in a similar manner.

The distillation is carried out at substantially lower pressure than the selective hydrogenation reaction, and the distillation pressure is generally about 1.1 to 10 bar. Thus, once the volatile components of the feed stream enter the distillation column 9, it will at least partially evaporate. Typically, about 50 to 100% of dissolved hydrogen will evaporate upon ingress into the column.

When preparing the TAME, the temperature at the top of the distillation column is about 40 to 70 ° C., typically about 50 to 60 ° C., and the temperature at the column bottom is about 100 to 150 ° C., typically about 120 to 130 ° C. to be.

As mentioned above, according to the invention, the distillation column of the reactive distillation apparatus is operated in such a way that the alkanol is heavier than hydrocarbons at the top of the distillation column. Thus, alkanols not bound to hydrocarbons in the form of azeotrope will attempt to flow downward in the column. At the same time, the vapor-liquid equilibrium of C ₅ and high molecular weight hydrocarbons with alkanols at the bottom of the column is such that alkanols are kept at lighter levels than hydrocarbons. As a result, alkanols flow upward from the bottom of the column. As such, the alkanol will circulate in the distillation system between the top and bottom of the column. By inducing a reaction bed in the distillation column, or by transporting the substream from the column through the reaction bed in the subreactor, an alkanol consuming reaction may occur that will remove the alkanol from the system.

Alkanols, in particular methanol and ethanol, form an azeotrope with the hydrocarbons of the feedstock. The higher the molecular weight of the hydrocarbon, the higher the alkanol concentration of the hydrocarbon-alkanol azeotrope mixture. According to the invention, in order to minimize the amount of unreacted alkanol that is removed from the distillation process, in fact only C ₄ -hydrocarbon-alkanol azeotrope mixtures are taken as the ceiling product. These azeotropic mixtures are the lowest molecular weight hydrocarbon-alkanol azeotropic mixtures and have the lowest alkanol concentrations.

According to the present invention, the amount of unreacted alkanol can be adjusted by adjusting the amount of C ₄ hydrocarbons in the feed to correlate with the amount of alkanol. The smaller the amount of C ₄ hydrocarbons in the feed, the less distillate can be removed and the less alkanol removed from the process. By increasing the amount of C ₄ hydrocarbons in the feed, the distillate flow rate can be increased without any variation in the relative amount of free unreacted alkanol in the ceiling product. Thus, if desired, C ₄ hydrocarbons (or even C ₃ hydrocarbons) are intentionally added to the process so that the intended effect can be achieved.

In the main fractionator (9), low molecular weight hydrocarbons and excess hydrogen are taken as vapor distillates together with small amounts of removed C ₄ hydrocarbons and methanol. This stream is used to regulate the pressure of the main fractionator. In the absence of a vapor stream, the distillate stream is generally very small relative to the reflux stream (reflux ratio generally exceeds 100), which will make it difficult to control the pressure of the main fractionator. It is not possible to use distillate streams for any quick control purposes. However, according to the present invention, distillation can be operated under stable conditions by adjusting the amount of gaseous distillate by means of a pressure valve and keeping the pressure at least nearly constant. The hydrogen and methane released together with any hydrocarbon azeotrope and other hydrocarbon atoms can be recycled to recover hydrogen. Preferably, the compound is delivered to the flue gas system.

When carrying out the process according to the invention, the alkanol concentration of the bottom product of the column can easily be reduced to a value as small as desired. In the case of methanol, it is possible to reduce the concentration of methanol in the bottom product to less than 100 ppm. The amount of alkanol in the distillate will only correspond to the amount bound by the azeotropic mixture. The composition of the azeotrope and the amount of alkanol removed are thus dependent on the hydrocarbon composition and distillation operating pressure of the ceiling product. Referring to an example based on the production of TAME: If C ₄ hydrocarbons account for the majority (> 90%) of the ceiling product, there will be about 0.1 to 5.0% by weight of methanol, depending on the distillation pressure and the amount of C ₅ hydrocarbons. . The more C ₅ hydrocarbons are included in the ceiling product, the more methanol will be removed with it (less than 90% by weight C ₄ hydrocarbon may be present in the ceiling product).

The etherification reaction described above is preferably carried out in the presence of a conventional cation exchange resin. However, other kinds of zeolites can also be used as etherification catalysts. As such, the resin may contain sulfonic acid groups, which may be obtained by sulfonation after polymerization and copolymerization of aromatic vinyl compounds. Examples of aromatic vinyl-based compounds suitable for preparing polymers of copolymers are styrene, vinyl toluene, vinyl naphthalene, vinyl ethylbenzene, methyl styrene, vinyl chlorobenzene and vinyl xylene. Acid cation exchange resins typically contain 1.3 to 1.9 sulfonic acid groups per aromatic nucleus. Preferred resins are based on copolymers of aromatic monovinyl compounds and aromatic polyvinyl compounds, in particular divinyl compounds, wherein the content of polyvinyl benzene is about 1 to 20% by weight of the copolymer. The ion exchange resins preferably have a granule size of about 0.15 to 1 mm.

In addition to the above resins, perfluorosulfonic acid resins which are copolymers of sulfonyl fluorovinyl ethyl and fluorocarbons can be used.

According to a preferred embodiment, the delivery location from the column to the subreactor is chosen in such a way that the vapor-liquid equilibrium ratio (K-value) of the alkanol is less than 1 in the (theoretical) tray above the delivery tray. If TAME is produced, the method can be carried out in such a way that the K-value of methanol is less than 1 in the tray above the delivery tray. In the TAEE method, there is generally one or two (theoretical) trays with a K-value less than 1 for ethanol between the delivery tray and the first tray above the delivery tray. The reaction product containing alkanol is returned to the column from the side reactor, which is fed to a tray having an alkanol K-value greater than one. As a result, alkanols are more hatched in the vapor phase than in hydrocarbons. The substream accounts for 40 to 90%, typically about 60 to 70%, of the total liquid flow rate in the column. The use of side reactors is preferred because, for example, the conditions prevailing in the distillation column can be influenced by varying the delivery location of the substream and by feeding a larger amount of alkanol to the reaction bed.

The present invention can also be applied to conventional catalytic distillation reactors. This reactor is operated in the same way as the subreactor process. The only difference is that the alkanol depletion reaction takes place in the column.

The invention is preferably carried out in conjunction with the MTBE, ETBE, TAME and TAEE processes.

In connection with the TAME process, the obtained ceiling product can be provided to the MTBE apparatus. Since the MTBE device contains some impurities (C ₅ hydrocarbons as far as the MTBE process is concerned), the ceiling product can be introduced into the feed of the MTBE device (which means that the C ₅ hydrocarbons remain in the MTBE product) or It can be introduced into the methanol washing apparatus of the MTBE apparatus. In the latter case, the C ₅ hydrocarbons are terminated with the raffinate stream of the MTBE apparatus, which contains mainly C ₄ hydrocarbons.

Alternatively, because the ceiling product of distillation contains only a small amount of methanol and the ceiling product is very small relative to the feed, the ceiling product by distillation may combine with the bottom product of the distillation to form a gasoline component. If necessary, the mixture is further processed. However, according to a preferred embodiment of the present invention, the C ₄ hydrocarbon content of the feed is intentionally kept small so that the mixture of the ceiling product and the bottom product can be used as a motor fuel component by itself.

Significant advantages are achieved by the present invention.

As described above, selective hydrogenation of the olefinic hydrocarbon feedstock is achieved by the present invention. The hydrogenation reaction will reduce the amount of diene to less than 0.1% (even less than 0.05%) and will also cause isomerization of the olefins, thereby increasing the amount of reactive olefins. When the hydrogenated product stream enters the distillation column after the preliminary reactor, the excess hydrogen and methane dispersed and dissolved in the hydrocarbon liquid phase and, if present, other volatile compounds will evaporate. The resulting gas phase can be used to control distillation.

At the same time, if excess hydrogen is delivered to the distillation column, no hydrogen separation system is required. If a separation device is used to separate excess hydrogen and low molecular weight hydrocarbons, some amount of low molecular weight hydrocarbons will still be present in the feed. Even when an overall condensation control scheme is used, non-condensable hydrocarbons tend to accumulate in the ceiling vessel. Therefore, these gases will sometimes have to be purged from the system, which affects pressure regulation and makes it difficult to maintain a stable operating pressure in the main fractionator. Moreover, if the low molecular weight hydrocarbon is separated immediately after selective hydrogenation, the pressure must be lowered. This means that since the hydrogenated product must be pumped into the etherification reactor, an extra pump is required after separation. Additional pumps increase the investment and operating costs.

Another disadvantage avoided by the present invention is the unwanted loss of reactive components during the separation operation.

The following examples will make the invention clear.

Example 1

Hydrogenation apparatus comprising a feed vessel, flow control means and a fixed bed reactor apparatus having a feed composed of 55 t FCC light gasoline mixed with 3 tonnes (t) of pyrolysis gasoline cut, as shown in FIG. Selectively hydrogenated. The hydrogen feed was 20 kg / h and the gas contained about 20 kg / h of other low molecular weight inert hydrocarbons, mainly methane. The pressure of hydrogenation was 25 bar and the temperature was 90 ° C. The WHSV used was 6. As catalyst, a Pd / Al ₂ O ₃ hydrogenation catalyst having 0.25 wt% palladium, supplied by BASF under the trade name H 0-22, was used.

In a mash feed of 20 kg H ₂ , 16 kg reacted, meaning that the H ₂ excess is 20% during the hydrogenation reaction.

Table 1 shows the results of the selective hydrogenation reaction.

Table 1 : Selective Hydrogenation of Olefinic Hydrocarbon Feeds

It can be seen from Table 1 above that the amount of diolefin is rapidly reduced to 0.05% and furthermore, the catalyst can isomerize some non-reactive components into reactive components.

Example 2

A feed consisting of 55 t FCC light gasoline mixed with 3 tons (t) of pyrolysis gasoline cut was selectively hydrogenated as described in Example 1 and then etherified in the etherification unit shown in FIG. .

Table 2 shows the component flow rates before and after the selective hydrogenation operation. In addition, the composition of the etherification products is described.

Table 2. Flow Rates of Hydrocarbon Feeds Before and After Selective Hydrogenation and Etherization

The composition of the distillate separated continuously during the operation of the distillation column is as follows: 20 mol% (1 wt%) of H ₂ , 13 mol% (5 wt%) of methane, 4 mol% (3 wt%) of methanol And 63 mol% (91 wt%) C ₄ -hydrocarbons. Many other low molecular weight components flow with hydrogen.

At a pressure of 500 kPa (absolute), the flow rate of gaseous distillate is about 41 m ³ / h. With this gas flow rate, the pressure in the distillation column is easily controlled. All of the above-mentioned components have a state similar to the ideal gas. This means that their relative volume corresponds to the molar fraction, ie hydrogen constitutes 20% of the total volume of the distillate. Together with low molecular weight hydrocarbons, the gas used for selective hydrogenation and carried to the distillation means for control of the distillation composition constitutes one third of the gaseous distillate. This additional amount of volatile components in the gas phase greatly improves the controllability of the distillation operation.

The invention will be explained in more detail by the accompanying drawings.

1 is a schematic illustration of a TAME process according to the present invention comprising an FCC gasoline splitter, an optional hydrogenation unit and an etherification unit.

2 is a view schematically showing a selective hydrogenation apparatus used in the present invention.

3 schematically shows an etherification apparatus comprising three pre-reactors, a product separation column and two side reactors.

Explanation of symbols on the main parts of the drawings

1: FCC gasoline splitter 2: Selective hydrogenation device

3: etherification apparatus 4: first hydrogenation reactor

5: second hydrogenation reactor 6, 7, 8: preliminary reactor

9: distillation column 10: condenser

11: reflux drum 12, 28 pump

13, 14: reactor

16, 17, 18, 19, 24, 25, 26: heat exchanger

20: reboiler

Claims (14)

In the presence of a catalyst, optionally a part or all of the polyunsaturated hydrocarbons is contacted by contacting a feedstock containing 4 to 7 carbon atoms, and a feedstock containing monounsaturated and polyunsaturated hydrocarbons with excess hydrogen to the polyunsaturated hydrocarbons. Hydrogenation to produce a modified hydrocarbon feedstock in which the unsaturated hydrocarbons consist primarily of monounsaturated compounds, Supplying some or all of the unreacted hydrogen and the modified hydrocarbon feedstock to a catalytic distillation reactor system comprising at least one distillation column, Reacting the C _4-7 isoolefin in the feedstock with an alkanol in the presence of a catalyst to produce a tertiary alkyl ether, Distilling the reaction mixture in a distillation column, Recovering virtually all unreacted hydrocarbons with alkyl ether and bottom product of distillation, Draining the ceiling product of distillation, which mainly contains an azeotropic mixture of C ₄ hydrocarbons and alkanols, and -Keeping the pressure of the distillation column substantially constant by controlling the amount of gaseous distillate containing hydrogen in the column, thereby producing a tertiary alkyl ether. The process according to claim 1, wherein 1.001 to 10 times molar excess of hydrogen is used for the hydrogenation of the hydrocarbon feedstock when calculated based on the polyunsaturated hydrocarbon of the feedstock. The process of claim 1, wherein the hydrogenation reaction is carried out at a pressure of 10 to 120 bar dissolving at least 80% hydrogen in the hydrocarbon feedstock during the hydrogenation reaction. 4. A process according to claim 3, wherein the hydrogenation reaction is carried out at a pressure of 15 to 115 bar. 4. A process according to claim 3, wherein the distillation is carried out at a substantially lower temperature than the hydrogenation reaction so that the hydrogen remains in gas phase during the distillation. The process of claim 1 wherein the hydrogenation reaction is carried out in the presence of a catalyst selected from the group consisting of supported noble metal catalysts and supported nickel catalysts. 7. A process according to claim 6, wherein the hydrogenation is carried out in the presence of a catalyst comprising palladium on alumina of from 0.2 to 5% by weight, when the concentration of palladium is calculated based on the total weight of the catalyst. The method of claim 1 wherein the temperature of the hydrogenation reaction is increased from 40 ° C. to 130 ° C. during operation to offset catalyst contamination. The process of claim 1, wherein the hydrogenation reaction is carried out in two or more successive reactors. The process according to claim 1, wherein the hydrogenation reaction is carried out using pure hydrogen gas. The process according to claim 1, wherein the hydrogenation reaction is carried out using hydrogen gas containing from 10 to 70% by weight of volatile low molecular weight paraffins. The process according to claim 1, wherein at least 5 mole% of excess hydrogen is supplied to the hydrogenation process, the pressure of the hydrogenation process is maintained at 15 bar or more, and the distillation pressure is maintained at 10 bar or less. The method of claim 1 wherein the alkanol comprises methanol or ethanol. The process according to claim 1, wherein the amount of distillate withdrawn from distillation corresponds to the amount of C ₄ hydrocarbons present in the feed.