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    • 专利摘要:
    The invention relates to a process for the preparation of ATBE and a C4 hydrocarbon stream comprising 1-butene and less than 2000 ppm by mass of isobutene based on the 1-butene, from a C4 hydrocarbon mixture (1) containing at least isobutene and 1 Containing by reacting an alcohol ROH with R = methyl or ethyl (2) with the present in the C4 hydrocarbon mixture isobutene on an acidic catalyst in an apparatus, the at least one pre-reactor (3), a reactive distillation column (5), at least a reaction zone (5b) and at least one distillation zone, wherein a distillation zone (5a) is arranged below the reaction zone, and at least one side reactor (10), wherein at least a portion of the internal liquid stream of the reactive distillation column as feed (9) driven into the side reactor is, the reaction mixture (11) is returned from the side reactor in the reactive distillation column, which is the pre-reactor leaving reaction mixture (4) is introduced below the reaction zone in the reactive distillation column, and wherein as the top product (7) of the reactive distillation column
    公开号:BE1018143A5
    申请号:E2007/0578
    申请日:2007-12-05
    公开日:2010-06-01
    发明作者:Armin Rix；Jochen Praefke；Fernandez Silvia Santiago；Ditk Roettger；Wilfried Bueschken
    申请人:Oxeno Olefinchemie Gmbh；
    IPC主号:
    专利说明:

    Process for the preparation of alkyl tert-butyl ether and high purity raffinate II
    The invention relates to a process for the preparation of high purity raffinate II (a C4 hydrocarbon mixture) having a low isobutene content, which is particularly suitable for the production of 1-butene, and alkyl tert-butyl ether (ATBE), in particular methyl tert. Butyl ether (MTBE) or ethyl tert-butyl ether (ETBE).
    Isobutene-free butene mixtures are suitable for the preparation of high-purity 1-butene and / or for the preparation of butene oligomers with low degrees of branching. ETBE and MTBE can be used as components of fuels for gasoline engines to increase the octane number. Due to the increasing availability of ethanol from renewable raw materials and the mandatory blending of biofuels, the demand for ETBE as a fuel additive is increasing.
    ATBE and linear butenes are recovered from C4-01efmgemischen, for example, the C4 cut from steam crackers or FCC units. These mixtures consist essentially of butadiene, the monoolefins isobutene, 1-butene and the two 2-butenes and the saturated hydrocarbons isobutane and n-butane. Conventional work-up procedures for such C4 cuts carried out worldwide comprise the following steps: First, most of the butadiene is removed. Butadiene can be well marketed or there is a self-consumption, it is z. B. separated by extraction or extraction distillation. In the other case, it is hydrogenated selectively to linear butenes up to concentrations of less than 1%. In both cases, a coal / hydrogen mixture remains (corresponding to raffinate I or hydrogenated crack), which contains the olefins (isobutene, 1-butene and 2-butenes) in addition to the saturated hydrocarbons (n-butane and isobutane). A possible way to remove the isobutene from this mixture is the reaction with methanol or ethanol to MTBE or ETBE. The saturated hydrocarbons, linear butenes and optionally an isobutene residue remain behind. The C4 mixture obtained after removal of the butadiene and isobutene is referred to as raffinate Π.
    Depending on the intended use of raffinate II, different isobutene contents are permissible.
    In addition to the linear olefins, raffinate II may contain relatively large amounts of isobutene, if this C4 mixture is reacted, for example, with acidic catalysts to give branched C4-oligomers, in particular Cs and C12-oligomers. This mixture gives after hydrogenation a hochoktanige gasoline fuel component.
    When raffinate II is used to prepare oligomers with low isoindices, i. H. low degree of branching, the content of isobutene must be very low, preferably less than 1% by mass. Virtually isobutene-free raffinate II is required if pure 1-butene is to be obtained from this raffinate II. The concentration of isobutene in the raffinate II should then not exceed those values that allow a workup to a 1-butene with a concentration of 2000 wppm. Depending on the work-up procedure, therefore, a value of 2000 wppm, preferably 1000 wppm and particularly preferably 600 wppm should not be exceeded. Since the boiling temperature difference between isobutene and 1-butene is only 0.6 ° C, an economical separation by distillation of the two components is not possible. In this case, isobutene has to be almost completely converted in the ATBE synthesis.
    The production of ETBE from isobutene-containing C4-hydrocarbon mixtures such as, for example, raffinate I or hydrogenated crack C4 with ethanol in the presence of acidic catalysts is analogous to the MTBE synthesis. The production of the ETBE is more difficult than the MTBE production due to the less favorable equilibrium position. Ethanol usually contains water, which must be largely removed, as it is the
    Reaction rate of ETBE formation and by-products, such as tert-butyl alcohol forms. Bio-alcohol contains additionally fusel oils (mainly C4- and C5-alcohols). Therefore, it is not easy to produce ETBE and a pure, virtually isobutene-free raffinate II. The preparation of MTBE and high-purity raffinate II is described, for example, in DE 101 02 082.
    For the production of ETBE from isobutene-containing C4 hydrocarbon streams such as, for example, raffinate I or hydrogenated crack C4 with ethanol, acid ion exchange resins (sulfonic acid groups) are frequently used as heterogeneous catalysts in the art. The
    Implementation takes place in one or more reactors connected in series, the catalyst preferably being arranged in a fixed bed. The result is a product in which ethanol, isobutene and ETBE are in equilibrium. In each reactor, the equilibrium conversion sets in according to the reaction conditions (temperature, ethanol excess, etc.). This mixture may then be separated by distillation into a bottom fraction containing ETBE and a top fraction containing C4 hydrocarbons and ethanol. After separation of the azeotrope-bound ethanol, the C4 hydrocarbon stream thus produced is not suitable for the production of 1-butene because of its high isobutene content.
    In order to achieve the highest possible isobutene conversion in the production of ETBE, the reaction of isobutene-containing C4-hydrocarbon mixtures with ethanol in several reaction stages, - after each reaction step, the ETBE formed is separated, or can be carried out in one stage using a reactive distillation column ,
    EP 0 071 032 describes a two-step process for the preparation of ETBE and a low isobutene C4 stream. In the first reactor, the isobutene-containing hydrocarbon stream is reacted with a superstoichiometric amount of ethanol and in a second reactor with a substoichiometric amount of ethanol. In the first reactor, ethanol is reacted with isobutene-depleted recycle streams. The effluent is separated in a first distillation column into a distillate consisting mainly of C4 hydrocarbons and a bottoms product consisting mainly of ETBE and ethanol. The isobutene poor distillate (product) is partially recycled to the first reactor. The bottom product of the first column is reacted with the distillate of the third column and feed C4 hydrocarbon mixture in the second reactor. The discharge of the second reactor is separated in a second column into a distillate, which is passed completely into the first reactor, and a bottom product consisting mainly of ETBE and ethanol. The bottom product of the second column is separated in a third column into ETBE and a distillate, which is passed into the second reactor. With this process, C4 hydrocarbon streams having a residual isobutene content of 0.3 mass% (3000 mass ppm or wppm) can be recovered.
    RU 2 167 143 describes a process for the preparation of ETBE. This process comprises two reaction stages and two distillation stages. The first reaction stage consists of one reactor or two reactors operated in a straight pass or in a loop mode. In it, an isobutene-containing hydrocarbon stream is reacted with ethanol on an acidic ion exchange resin to form a mixture containing ETBE. The effluent from the last reactor of this stage is separated in the first distillation column into ETBE (target product) and a distillate containing mainly isobutene and other Cr hydrocarbons. The distillate is reacted in a second reaction stage, which consists of one reactor or two reactors, with the addition of ethanol. This reaction mixture is separated in a second distillation column into a still isobutene-containing C4 hydrocarbon stream and an ethanol-containing ETBE, which is returned to the first reaction stage. The main disadvantage of this process is the relatively low isobutene conversion. According to Example 7 of this patent, the total isobutene conversion after the second stage is only 89%.
    The unpublished document DE 10 2005 062 722.6 describes a two-stage process for the preparation of ETBE from industrial mixtures of C 4 -hydrocarbons I which contain at least 1-butene, isobutene, n-butane and 2-butenes. The claimed process comprises the following process steps: a) reaction of parts of the isobutene contained in the technical mixture with ethanol in the presence of an acidic catalyst to ETBE, b) separation of the unreacted C4 hydrocarbons ΠΙ from the discharge of the stage a) by thermal separation processes c) distillative separation of the C4 hydrocarbons ΠΙ in a at least 1-butene and isobutene-containing fraction IV and a nearly isobutene-free, at least 2-butenes and n-butane containing fraction V, d) reaction the isobutene contained in fraction IV with ethanol VI in the presence of acidic catalysts to ETBE and e) separation of the unreacted C4 hydrocarbons VIII from the effluent of step d) to obtain a fraction VII containing ETBE.
    This process has the advantage that isobutene conversions of more than 99% can be achieved. However, it has the disadvantage of high complexity.
    US Pat. Nos. 5,368,691, 5,590,361 and 6,107,526 describe the preparation of ETBE by reacting isobutene-containing C4 hydrocarbon streams with ethanol in an apparatus consisting of a pre-reactor and a reactive distillation column. In these methods, the focus is on the recovery of an ethanol-poor ETBE. Isobutene conversions or residual isobutene contents in the distillate are not disclosed.
    EP 0 820 794 claims a process for preparing ETBE by reacting ethanol with an isobutene-containing C 4 -hydrocarbon mixture in the presence of an acidic catalyst. The reaction takes place in a plant which consists of a pre-reactor, a distillation column, a reactive distillation column and a further reactor which is used as intermediate reactor or as side reactor of one of the two columns. The educts ethanol and C4 hydrocarbon mixture are reacted in the prereactor. The resulting reaction mixture is separated in the distillation column into a bottom product containing mainly ETBE and a top product with the unreacted C4 hydrocarbons. This mixture is passed into the lower part of the reactive distillation column below the catalyst zone. In the reactive distillation, a further part of the isobutene is reacted. As its top product, an ethanol-containing isobutene-depleted C4 distillate is stripped off. The bottoms product is pumped as reflux to the top of the first column. Depending on the process variant, the further reactor may be located at different points in the system. In a process execution he lies between the two columns. It is flowed through with or without addition of ethanol from the bottom product of the reactive distillation column. In another variant, the first column has a side reactor and in a further variant, the reactive distillation column. The side reactors are operated with the addition of ethanol. All these process variants have the main disadvantage that the total isobutene conversion is below 98%.
    WO 94/08927 describes a method for obtaining ETBE by reacting a
    Isobutene-containing C4 hydrocarbon mixtures with ethanol in a plant consisting of a reactor and a reactive distillation column. In this case, in the pre-reactor, the feed hydrocarbon mixture is reacted with the reflux from the catalyst zone of the reactive distillation column containing ethanol, ETBE and hydrocarbons. The reaction mixture leaving the reactor is introduced into the lower part of the reactive distillation column below the catalyst zone. Ethanol is fed exclusively into the upper region of the reactive distillation column above the catalyst zone. The bottom product obtained is a mixture consisting mainly of ETBE, and the top product is an ethanol-containing C4 hydrocarbon mixture. On page 11, line 24 of this document it is said that the isobutene is essentially completely reacted. However, it is not possible to deduce any information about the actual isobutene content from the text because no information is given on the compositions of the starting materials, the bottom product and the distillate and the ratio of isobutene used to ethanol used.
    WO 93/19031 describes a process for the preparation of ETBE by reacting an isobutene-containing C 4 -hydrocarbon mixture with ethanol. The reaction takes place in an apparatus which consists of at least one pre-reactor and a reactive distillation column with side reactor which contains at least 30% of the total amount of catalyst. The major part of the turnover is carried out in the side reactor. The feed hydrocarbon mixture is reacted with ethanol in the prereactor. The reaction mixture leaving the prereactor is introduced into the lower part of a reactive distillation column below the catalyst zone. From the column, a portion of the contents is withdrawn and this optionally reacted with the addition of further ethanol in a side reactor. The discharge of the side reactor is returned to the column at a lower point than its inflow was removed. The top product is an ethanol-containing C4 stream and withdrawn as the bottom product, a stream consisting mainly of ETBE. A disadvantage of this process is that no complete isobutene conversion is achieved. In Example 3, the distillate (see table on page 21) still has a content of isobutene of 6.9%.
    The virtually complete separation of isobutene from a C4 hydrocarbon stream, the
    Isobutene and 1-butene contains, by reacting isobutene with ethanol to ETBE can, as already described, technically done by a two-step process. However, this has the disadvantage of a high capital investment and high operating costs. By contrast, one-step processes are characterized by lower capital investment and lower operating costs, but have the disadvantage that almost complete isobutene conversion is not guaranteed. It was therefore an object to develop a process that has the advantages of a two-stage process and that of a one-stage process.
    Surprisingly, it has now been found that from a isobutene and 1-butene containing C hydrocarbon stream ETBE and a hydrocarbon stream with less than 2000 ppm by mass (wppm) isobutene can be prepared in one stage, if the reaction of the isobutene present in the C4 hydrocarbon stream with ethanol is carried out in an apparatus which consists of at least one pre-reactor and a reactive distillation column with side reactor. From such a hydrocarbon stream comprising 1-butene and less than 2000 (wppm), preferably less than 1000 wppm and preferably less than 600 wppm of isobutene, 1-butene with less than 2000 wppm of isobutene can be readily prepared. That a higher isobutene conversion is achieved with a reactive distillation column with side reactor than with a reactive distillation without side reactor was particularly surprising because a model calculation for the production of ETBE by the reaction of ethanol with an isobutene-n-butane mixture the opposite (BH Bisowamo, Y.C. Tian, Mo Tade, Application of side reactors on ETBE reactive distillation, Chemical Engineering Journal 99 (2004), pages 35-48).
    The present invention therefore relates to a process for the preparation of ATBE and a C4 hydrocarbon stream which comprises 1-butene and less than 2000 ppm by mass of isobutene, based on the 1-butene present in the C4 hydrocarbon stream, from a C4-hydrocarbon mixture (US Pat. 1) containing at least isobutene and 1-butene, by reacting an alcohol ROH with R = alkyl having 1 to 5 carbon atoms, in particular methyl or ethyl (2) with the isobutene present in the C4 hydrocarbon mixture on an acidic catalyst in an apparatus, the at least one pre-reactor (3) and a reactive distillation column (5), the at least one reaction zone (5b) and at least one
    Destillation zone has, wherein a distillation zone (5 a) is arranged below the reaction zone, and at least one side reactor (10), wherein at least a portion of the internal liquid stream of the reactive distillation column is run as feed (9) in a side reactor, the reaction mixture (11 ) is returned from the side reactor in the reactive distillation, the reaction mixture leaving the pre-reactor (4) is introduced below the reaction zone in the reactive distillation, as top product (7) an alcoholic Cr hydrocarbon stream and the bottom product (6) at least 50 wt .-% ATBE-containing stream which is characterized in that in the side reactor, a catalyst volume is used which corresponds to 10 to 160% of the catalyst volume in the reactive distillation column.
    In addition, the present invention is a composition containing at least 25% by mass of 1-butene, 0.5 to 7% by mass of alcohol ROH with R = methyl or ethyl and between 10 and 2000 wppm isobutene.
    The inventive method has a number of advantages. It requires less capital input and lower operating costs compared to two-stage processes. With the aid of the process according to the invention, a distillate can be obtained with less than 2000 wppm, preferably less than 1000 wppm and preferably less than 600 wppm of isobutene relative to the C 4 -hydrocarbons, which is suitable for the production of 1-butene with less than 2000 ppm by weight. ppm isobutene is suitable. Furthermore, only a small amount of 1-butene is lost by isomerization to 2-butenes or other side reactions such as oligomerization. Another advantage is that the catalyst in the side reactor can be changed more easily than the catalyst in the packing of the reactive distillation column. This can also lead to a higher availability of the entire system when the catalyst is changed in the side reactor in a reduced load phase and thus a complete shutdown of the process / the system is not necessary. The use of the side reactor also requires fewer catalyst packings in the reactive distillation column, resulting in longer service lives. From the smaller number of necessary catalyst packages also results in a lower overall height for the reactive distillation column, whereby the
    Construction of Reaktivdestiliationskolonne and auxiliary equipment is less expensive.
    The inventive method also has the advantage that it is particularly well suited to retrofit existing plants for the production of MTBE in conjunction with the recovery of 1-butene on the production of ETBE, without quality losses have to be taken in winning 1-butene ,
    The method according to the invention will be described below by way of example, without the invention, the scope of which results from the claims and the entire description, being restricted thereto. The claims themselves belong to the disclosure of the present invention. If regions or preferred ranges are specified in the following text, then all theoretically possible subareas lying in these ranges should also belong to the disclosure content of the present invention, without these having been explicitly named for reasons of clarity.
    The inventive method for producing ATBE and a C4 hydrocarbon stream comprising 1-butene and less than 2000 ppm by mass of isobutene based on the 1-butene, from a C4 hydrocarbon mixture (1) containing at least isobutene, in particular in a Concentration of 2000 wppm based on the 1-butene, and containing 1-butene, by reacting an alcohol ROH with R = alkyl having 1 to 5 carbon atoms, in particular with R = methyl or ethyl (2) with the present in the C4 hydrocarbon mixture isobutene on an acidic catalyst in an apparatus comprising at least one pre-reactor (3), a reactive distillation column (5) having at least one reaction zone (5b) and at least one distillation zone, wherein a distillation zone (5a) is located below the reaction zone, and at least one Side reactor (10), wherein at least a portion of the internal liquid stream of the Reaktivdestiliationskolonne driven as feed (9) in the side reactor is, the reaction mixture (11) is returned from the side reactor in the Reaktivdestiliationskolonne, the pre-reactor leaving the reaction mixture (4) is introduced below the reaction zone in the Reaktivdestiliationskolonne, and wherein as the top product (7) of the Reaktivdestiliationskolonne an alcoholic C4 hydrocarbon stream and as the bottom product (6) at least 50 mass% ATBE
    withdrawing stream is distinguished by the fact that in the side reactor, a catalyst volume is used, which corresponds to 10 to 160%, preferably 20 to 120%, preferably 25 to 100% and particularly preferably 30 to 75% of the catalyst volume in the reactive distillation column. It may be particularly advantageous if the process is carried out such that at most 25% by volume, preferably from 1 to 25% by volume, preferably from 2 to 20% by volume and more preferably from 3 to 15% by volume of the material present in the apparatus Total catalyst volume (sum of all catalyst volume in the pre-reactors, the reactive distillation column and the side reactor) is present in the side reactor.
    The erfmdungsgemäße 1-stage process has three process steps, in which a reaction of the isobutene with the alcohol takes place, namely a reaction in at least one prereactor, a reaction and separation in the reactive distillation and a reaction in the side reactor. These three steps are described in detail later.
    In all three process steps, an acidic catalyst is used. The catalysts used in the individual process steps may be the same or different. Even in a single process step, one or more catalysts can be used.
    The catalyst used in the process according to the invention is preferably a solid having acidic centers on its surface. Preferably, the catalyst used is soluble neither in one of the starting materials nor in the product mixture. The catalyst is preferably selected so that it does not release any acidic substances to the reaction mixture under reaction conditions, since otherwise this could lead to loss of yield.
    In the process according to the invention, preference is given to using catalysts which, under reaction conditions, catalyze the addition of alcohol to isobutene but scarcely the addition to linear butenes. Particular preference is given to using catalysts which have no or only very little activity under the reaction conditions with respect to the oligomerization of olefins and dialkyl ether formation. So a high
    Yield of 1-butene can be achieved with the method according to the invention, it may also be advantageous to use a catalyst having little or no activity with respect to the isomerization of 1-butene to the 2-butenes.
    As solid catalysts, for example, zeolites, acid-activated bentonites and / or clays, sulfonated zirconium oxides, montmorillonites or acidic ion exchange resins can be used. Particular preference is given to using acidic catalysts (acidic) ion exchange resins in the process according to the invention.
    A group of acidic catalysts preferably used in the process according to the invention are solid ion exchange resins, in particular those having sulfonic acid groups. Suitable ion exchange resins may be, for example, those prepared by sulfonation of phenol / aldehyde condensates or cooligomers of vinyl aromatic compounds. Examples of aromatic vinyl compounds for the preparation of the cooligomers are: styrene, vinyltoluene, vinylnaphthalene, vinylethylbenzene, methylstyrene, vinylchlorobenzene, vinylxylene and divinylbenzene. In particular, the co-oligomers formed by reacting styrene with divinylbenzene are used as precursors for the preparation of ion exchange resins having sulfonic acid groups. The resins can be prepared gel-like, macroporous or spongy.
    The properties of these resins, especially specific surface area, porosity, stability, swelling and exchange capacity, can be varied by the manufacturing process.
    In the process according to the invention, the ion exchange resins are used in their H form, wherein optionally a part of the H ions can be exchanged for other cations, in particular metal cations. Strongly acidic resins of the styrene-divinylbenzene type become u. a. sold under the following trade names: Duolite C20, Duolite C26, Amberlyst 15, Amberlyst 35, Amberlite IR-120, Amberlite 200, Dowex 50, Lewatit SPC 118, Lewatit SPC 108, K2611, K2621, OC 1501.
    The pore volume of the ion exchange resins preferably used is preferably from 0.3 to 0.9 ml / g, in particular from 0.5 to 0.9 ml / g. The grain size of the preferably used resins is preferably from 0.3 mm to 1.5 mm, in the packing of the reactive distillation column, in particular from 0.5 mm to 1.25 mm. The particle size distribution can be narrowed or further selected. Particular preference is given to using ion exchange resins having a very uniform particle size in the packing of the reactive distillation column. The capacity of the ion exchange resins preferably used, based on the delivery form, preferably from 0.7 to 2.0 eq / 1, in particular from 1.1 to 2.0 eq / 1, or preferably from 0.5 to 5.5 mol / kg, in particular from 0.8 to 5.5 mol / kg (The information on the capacity in mol / kg refers to the ion exchange resin dried to constant weight in the warm stream of nitrogen, eg at 105 ° C.).
    In the reaction zone of the reactive distillation, the same catalysts can be used as used in the pre- and side reactors. In the reactive distillation column, the catalyst can either be integrated in the packing, for example KataMax® (as described in EP 0 428 265), KataPak® (as described in EP 0 396 650 or DE 298 07 007.3 U1) or on shaped bodies (as described in US Pat in US 5,244,929). In the reactive distillation column, the acidic catalyst is particularly preferably used in the form of tissue packs which enclose the acidic catalyst.
    feedstocks
    As C 4 -hydrocarbon mixtures, it is possible to use in the process according to the invention all customary technical C 4 -hydrocarbon mixtures which comprise isobutene and 1-butene. Suitable isobutene and 1-butene-containing C4 streams are, for example, light petroleum fractions from refineries, C4 fractions from crackers (for example steamer, hydrocracker, catalracker), mixtures of Fischer-Tropsch syntheses, mixtures from the dehydrogenation of butanes, mixtures from skeletal isomerization linear butenes, mixtures formed by metathesis of olefins or mixtures of other technical processes. Some of these techniques are known in the literature, e.g. In K.Weissermel, H.J. Arpe, Industrial Organic Chemistry, Wiley-VCH, 5th Edition, 1998, pages 23-24; 65-99; 122-124.
    Preference is given to using C4 fractions from steam crackers, which are operated primarily for the production of ethene and propene and in which, for example, refinery gases, naphtha, gas oil, LPG (liquefied petroleum gas) and NGL (natural gas liquid) are used as raw materials, or Katcrackem. Depending on the cracking process, the by-produced C4 cuts contain varying amounts of isobutene. Other main constituents are 1,3-butadiene, 1-butene, c-2-butene, t-2-butene, n-butane and i-butane. Typical isobutene contents in the C4 fraction are from 18 to 35% by mass for steam cracking C4 fractions and 10 to 20% by mass for fluid catalytic cracking (FCC).
    For the process according to the invention, it may be advantageous to remove polyunsaturated hydrocarbons, such as 1,3-butadiene or alkynes, from the C 4 -hydrocarbon mixture. C 4 -hydrocarbon mixtures which comprise isobutene and 1-butene and which contain less than 1000 wppm of alkynes and / or less than 8000 ppm of 1,3-butadiene are particularly preferably used in the process according to the invention. Such mixtures (raffinate I or selectively hydrogenated crack C4) are a preferred starting material for the process according to the invention.
    The removal of polyunsaturated hydrocarbons from the C 4 -hydrocarbon mixture can be carried out by known processes, for example by extraction, extractive distillation or complex formation (cf., K.Weissermel, HJ Arpe, Industrielle Organische Chemie, Wiley-VCH, 5th edition, 1998, pages 119 to 121).
    An alternative to the separation of the polyunsaturated hydrocarbons is a selective chemical reaction. For example, 1,3-butadiene can be selectively hydrogenated to linear butenes, such as. As described in EP 0 523 482. Also, by selective reactions of the 1,3-butadiene, for example dimerization to cyclooctadiene, trimerization to cyclododecatriene, Polymérisations- or telomerization reactions, the 1,3-butadiene can be at least partially removed. When a crack C4 cut has been used as the feedstock, a hydrocarbon blend (eg, raffinate I or hydrogenated crack C4 (HCC4)) remains in all cases, mainly the saturated hydrocarbons, n-butane and isobutane and the olefins isobutene , 1-butene and 2-butenes and optionally radicals of 1,3-butadiene.
    In the process according to the invention, preference is given to catalytically selectively hydrogenating polyunsaturated hydrocarbons present in the C4 hydrocarbon streams in an additional purification stage which is preceded by reaction with alcohol. Particular preference is given to such a purification stage, if it can not be ruled out that the technical C 4 hydrocarbon streams used have polyunsaturated hydrocarbons in larger amounts than stated above.
    The polyunsaturated hydrocarbons are mainly 1,3-butadiene; 1,2-butadiene, butenine and 1-butyne are contained in a significantly smaller amount, if at all. The hydrogenation can take place in a single-stage or multistage hydrogenation process in the liquid phase on a palladium contact. In order to lower the content of 1,3-butadiene under preferably 1000 ppmw, the hydrogenation is carried out in the last stage with addition of a moderator, which increases the selectivity of the palladium contact. The moderator used is preferably carbon monoxide, which is added in a proportion of 0.05 to 100 ppm by mass (wppm). The content of polyunsaturated hydrocarbons in the feed to this stage should be less than 1%, preferably less than 0.5%. This type of selective hydrogenation of residual contents of 1,3-butadiene is known in the literature under the name SHP (selective hydrogénation process) (compare EP 0 081 041, petroleum, coal, natural gas, Petrochem 1986, 39, 73).
    If the isobutene-containing C4 streams contain more than 1% by mass of polyunsaturated hydrocarbons, such as 1,3-butadiene, they are preferably reacted in upstream hydrogenations. These hydrogenations are preferably carried out in the liquid phase on a palladium contact. Depending on the content of polyunsaturated hydrocarbons, the hydrogenation can be carried out in several stages. For the conversion of crack C4 from a steamer with a content of 1,3-butadiene of typically 38 to 45%, a two-stage version of the hydrogenation has proven itself. In this case, individual or all stages may be equipped with a partial product return. Concentrations of 1,3-butadiene less than 1% are thus obtainable in the discharge, so that further conversion into a selective hydrogenation (SHP) can take place.
    The technical hydrocarbon mixtures used in the process according to the invention, which have isobutene and 1-butene, preferably have the following compositions, wherein, depending on the content of unsaturated hydrocarbons, a hydrogenation or selective hydrogenation is carried out before the reaction with alcohol.
    The raffinate I or HCC4 is, among others, a preferably used isobutene-containing
    Hydrocarbon mixture in the context of this invention. Since plants for the processing of Cj-hydrocarbons are usually constructed as a strand (combination of several plants), it is possible, however, that the raffinate I or HCC4 prior to entering the process of the invention one or more other process stage (s). This process stage (s) can also be, for example, a method or method step (e) as have been described in the embodiments for method step a). C4 hydrocarbon mixtures which can be used in the process according to the invention can also be those obtained from processes according to the embodiments of process step a) and subsequent separation according to process step b). In particular, it is also possible to use those mixtures obtained in the preparation of tert-butanol (TBA) from isobutene after separation of the TBA. In this way, an individually adapted overall concept for processing with the corresponding product portfolio can be realized in each case.
    Typical processes which may be preceded by the processes according to the invention are water washing, cleaning processes in adsorbers, drying processes and distillations.
    By washing with water, hydrophilic components from the technical hydrocarbon mixture containing isobutene and 1-butene can be completely or partially removed, for example nitrogen components. Examples of nitrogen components are acetonitrile or N-methylpyrrolidone (which may be obtained, for example, from a 1,3-butadiene extractive distillation). Oxygen compounds (eg acetone from FCC) can also be partially removed by washing with water. The isobutene-containing hydrocarbon stream is saturated after a water wash with water. In order to avoid a two-phase reaction in the subsequent process steps in the reactor, the reaction temperature should be about 10 K above the temperature of the water wash.
    Adsorbers can be used to remove impurities. This can be advantageous, for example, if noble metal catalysts are used in one of the process steps. Often, nitrogen or sulfur compounds are removed via upstream adsorbers. Examples of adsorbents are aluminas, molecular sieves, zeolites, activated carbon, metals impregnated clays. Adsorbents are sold by various companies, for example the company Alcoa (Selexsorb®).
    Water optionally contained in the C4 hydrocarbon mixture containing isobutene and 1-butene, which may originate for example from a water wash, can be removed by known drying processes. Suitable methods are, for example, the distillative separation of the water as an azeotrope. In this case, an azeotrope with contained C4 hydrocarbons can often be utilized or entrainers can be added.
    Drying of the C 4 -hydrocarbon mixture can be advantageous for a variety of reasons, for example to reduce the formation of tert-butanol or to avoid technical problems due to separation of water or to avoid ice formation at low temperatures (eg during intermediate storage). ,
    Distillation steps can be used, for example, to separate impurities from the technical hydrocarbon mixture (for example low boilers such as C3 hydrocarbons, high boilers such as Cs hydrocarbons) or to obtain fractions having different isobutene concentrations. By direct distillation of the technical KohlenwasserstofFgemische example, a separation in a 2-butenes and n-butane impoverished, isobutene richer fraction can be achieved.
    Depending on the composition of the technical hydrocarbon mixture to be used, the technical hydrocarbon mixture can thus be reacted directly with the alcohol in the process according to the invention or can only be used after a pretreatment by one or more of the abovementioned processes.
    As a second feedstock, in the process according to the invention, highly pure alcohol, pure alcohol or alcohol can be used, which preferably has small amounts of impurities. Preferably, the purity of the alcohol used, expressed in% by mass of alcohol, is greater than 90%, more preferably greater than 95%, and most preferably equal to or greater than 99%. Alcohols can be alkanoic acid with 1 to 5
    Carbon atoms, in particular methanol or ethanol can be used. Ethanol is preferably used in the process according to the invention. The purity of the ethanol used, stated in% by mass of ethanol, is preferably greater than 90%, particularly preferably greater than 95% and very particularly preferably equal to or greater than 99%. Ethanol, which has a purity of greater than or equal to 99% by mass, can, for. B. be bioethanol. The content of water is preferably below 3% by mass, more preferably below 1% by mass, most preferably below 0.5% by mass. The use of alcohol can z. B. dried by azeotropic distillation and / or membrane process and / or entrainer distillation.
    Particular preference is given to using denatured ethanol as ethanol. Very particular preference is given to using as ethanol an ethanol which is in the form of a denaturant ETBE, preferably in a concentration of 0 to 5% by mass, preferably of 0.005 to 1% by mass, particularly preferably of 0.05 to 1% by mass and more preferably from 0.01 to 0.2 mass%. In Germany, ethanol is particularly preferably used which has at least 0.085% by volume denaturant.
    pre-reaction
    The educts, ie alcohol and C 4 -hydrocarbon mixture comprising isobutene and 1-butene, are reacted in a first reaction with one another in the first process step of the process according to the invention. The pre-reaction can be carried out in one or more pre-reactor (s). The pre-reactors can be conventional fixed-bed reactors (tube-bundle reactors, adiabatic fixed-bed reactors, circulation reactors etc.), stirred tanks or a combination thereof. Preferably, the reaction conditions in at least one pre-reactor will be chosen so that the isobutene is reacted with the alcohol at the exit of the reaction mixture from the last pre-rector to or close to equilibrium with ATBE.
    The prereactor / pre-reactors can / can be operated at temperatures of preferably 10 to 160 ° C, preferably from 30 to 110 ° C. The pressure in the prereactor / pre-reactors is preferably 5 to 50 barabS. (bara), preferably 7.5 to 20 bara and more preferably from 8 to 13 bara.
    The molar ratio of alcohol to isobutene in the pre-reactors is preferably from 5 to 1 to 0.9 to 1, preferably from 2 to 1 to 1 to 1 and more preferably from 1.2 to 1 to 1 to 1. The ratio may be in the pre-reactors are adjusted by the first pre-reactor or each of the existing pre-reactors, an appropriate amount of alcohol is supplied.
    In a preferred embodiment, the reaction of the isobutene is carried out in at least two pre-reactors, wherein the first prereactor is operated as adiabatic fixed bed reactor with recycle (circulation reactor) and the following prereactor / the following prereactors as Festbettstufe / -n without recirculation / are. The ratio of recycled amount to Frischzulauf ^ hydrocarbons and alcohol, especially ethanol) is preferably 0.5 to 20 t / t, more preferably from 1 to 5 t / t and most preferably from 2 to 3 t / t. The circulation reactor preferably has an inlet temperature of 35 to 50 ° C and an outlet temperature of 50 to 70 ° C and is preferably operated at 10 to 13 bara. Since the thermodynamic equilibrium between alcohol / isobutene and ATBE, especially between ethanol / isobutene and ETBE at low temperature predominantly on the side of the ether, it is preferred to use the first of the reactors at a higher temperature (high reaction rate) than the following (utilization of the equilibrium position) to operate. Particularly preferred for the pre-reaction, a reactor system is used which has three series-connected pre-reactors, of which the first pre-reactor is operated as a circulation reactor and the two subsequent pre-reactors are operated in a straight pass. It may be advantageous if several, preferably two of these pre-reactor systems are present for the pre-reaction, so that in repair work, such. As catalyst change, on a prereactor of the Vorreaktorsysteme in the other prereactor system without interrupting the process (but with reduction of throughput) can be carried out further. When using pre-reactor systems of three reactors, the pre-reactors are preferably operated at a temperature of 30 to 80 ° C, preferably 40 to 75 ° C and a pressure of 5 to 20 bara, preferably 7 to 15 bara, preferably the temperature in the reactors is set lower from first to last reactor. The pre-reactor / pre-reactors downstream of the circulation reactor preferably have an inlet temperature of 30 to 50 ° C, preferably 30 to 40 ° C and an outlet temperature of 35 to 45 ° C, and are preferably also operated at 8 to 13 bara.
    Reaction and separation in the reactive distillation column
    The reaction mixture obtained from the first process step of the preliminary reaction is further treated in a reactive distillation. The reactive distillation is preferably carried out in a reactive distillation column which has at least one reaction zone and at least one distillation zone. In the reaction zone, both reaction and distillation preferably take place. The reactive distillation column used may have one or more reaction zones, preferably followed by a distillation zone above the uppermost and below the lowermost reaction zone. With particular preference, the reactive distillation column has a reaction zone, to which above and below each followed by a pure distillation zone.
    The reaction zone preferably has a separation capacity of 5 to 100, preferably from 20 to 50 theoretical plates. The distillation zone arranged below the lowest reaction zone preferably has a separation capacity of from 10 to 40, preferably from 15 to 35 and very particularly preferably from 15 to 25 theoretical plates. If a distillation zone is present above the uppermost reaction zone, it preferably has from 2 to 20, preferably from 5 to 15 and very particularly from 5 to 10 theoretical plates.
    The feed for the reaction product from the prereactor or prereactor system to the reactive distillation column is preferably carried out in the region of the distillation zone arranged below the lowermost reaction zone. The feed to the reactive distillation column preferably takes place in the range from 3 to 15, particularly preferably in the range from 6 to 12 theoretical separation stage of this distillation zone (counting of the separation stages from bottom to top).
    Should it be necessary or desired to introduce further alcohol into the reactive distillation column, this is preferably carried out in the region of the reaction zone.
    It may be advantageous if more alcohol is present in the feed of the reactive distillation column (reaction mixture from the pre-reaction) than is needed for the complete conversion of the remaining isobutene. The excess alcohol should preferably be at least sufficient to provide enough alcohol to form the alcohol and C4 hydrocarbon azeotrope. In order to achieve a conversion of isobutene of more than 99.5%, in the process according to the invention, moreover, preferably a higher alcohol excess, in particular excess ethanol, is used. The amount of alcohol required for this purpose can be fed into the reaction zone of the reactive distillation column and / or into the side reactor.
    The temperature of the feed from the pre-reactor or pre-reactor system to the reactive distillation column, regardless of its composition, of the reaction pressure in the column and the throughput preferably between 50 ° C and 80 ° C, preferably from 60 ° C to 75 ° C, particularly preferably just below or at, preferably below the boiling point of the feed mixture at column pressure.
    The reaction of the isobutene with alcohol to ATBE is carried out in the reactive distillation, depending on the pressure preferably in the temperature range from 40 ° C to 120 ° C, preferably from 60 ° C to 90 ° C, more preferably from 65 ° C to 85 ° C (temperature in the catalyst zone, the temperature in the bottom of the column can be significantly higher). If isobutene is reacted with ethanol to form ETBE in the process according to the invention, the reaction in the reactive distillation column preferably takes place in the temperature range from 40 ° C. to 120 ° C., preferably from 60 ° C. to 90 ° C., particularly preferably from 65 ° C., depending on the pressure up to 85 ° C.
    The reaction of isobutene with alcohol in the reactive distillation column is preferably carried out at pressures, measured at the top of the column, from 3 bara to 15 bara, preferably from 7 bara to 13 bara, and more preferably from 8 bara to 12 bara.
    The hydraulic load in the catalytic packing of the column is preferably 10% to 110%, preferably 20% to 70% of its flood point load. Hydraulic loading of a distillation column is understood to be the uniform flow stress of the column cross-section due to the ascending vapor mass flow and the returning liquid mass flow. The upper load limit indicates the maximum load of steam and return fluid, above which the separation effect due to entrainment or jamming of the return fluid by the rising vapor flow decreases. The lower load indicates the minimum load below which the separation effect decreases or breaks down as a result of irregular flow or idling of the column -z. As the soils (Vauck / Müller, "Basic Operations Chemical Process Engineering", p 626, VEB German publishing house for basic industries.). At the point of flooding, the shear stresses transferred from the gas to the liquid become so great that the entire amount of liquid is entrained in the form of drops with the gas or that phase inversion occurs in the column (J. Mackowiak, "Fluid Dynamics of Columns with Modern Packages and Packages for gas / liquid systems ", Otto Salle Verlagl991).
    In the process according to the invention, it may be advantageous if the reactive distillation column are operated with reflux ratios of less than 1.5, preferably with reflux ratios of 0.6 to 1.2, preferably between 0.7 and 1.1.
    As the top product of the reactive distillation column, an alcoholic, in particular ethanol-containing distillate can be withdrawn, which, based on the C4 hydrocarbons therein, a residual isobutene concentration of less than 2000 wppm, preferably less than 1000 wppm and more preferably less than 600 wppm. Preferably, the top product has a concentration of isobutene of less than 2000 wppm based on the 1-butene in the top product. From the top product can, for. B. by an extraction step, in particular by extraction with water, the alcohol contained are separated. From the product thus obtained, which is often referred to as raffinate Π, traces of butadiene can be removed by selective hydrogenation (SHP). This mixture can be separated by distillation into 1-butene, isobutane and a mixture of 2-butenes and n-butane or in 1-butene, 2-butene and n-butane.
    The top product C4 hydrocarbon stream can, for. B. with the described
    Methods, to a 1-butene, which has a content of isobutene based on the contained 1-butene of less than 2000 wppm, preferably less than 1500 wppm Auiweist, worked up. Thus produced (pure) 1-butene is a sought-after intermediate product. It is used, for example, as a comonomer in the production of polyethylene (LLDPE or HDPE) and of ethylene-propylene copolymers. It is also used as alkylating agent and is the starting material for the production of butan-2-ol, butene oxide, valeraldehyde.
    Another use of the inventively produced almost isobutene-ffeien, raffinate II is the production of n-butene oligomers, especially after the octol process.
    The hydrocarbons remaining after separation or reaction of the linear butenes from the raffinate II can optionally be worked up to form isobutane and n-butane after hydrogenation (CSP).
    The bottom product of the reactive distillation is a mixture which preferably contains at least 75% by mass, preferably at least 90% by mass and more preferably at least 95% by mass of ATBE. In addition, it contains some of the excess ethanol. This mixture can be used as a component for gasoline fuels. If desired, this mixture can be worked up to ATBE with a higher purity.
    Optionally, from the lower part of the reactive distillation column, preferably from the region of the lowermost distillation zone, a side stream rich in alcohol, in particular ethanol, can be withdrawn, which can be recycled as alcohol source or ethanol source, optionally after a further work-up into the apparatus, for example into the prereactor ,
    The reactive distillation column used in the process according to the invention is connected to a side reactor in which the third process step is carried out.
    The removal of the feed to the side reactor takes place in the reactive distillation column, preferably below the lowest reaction zone. Preferably, the removal of the
    Feed to the side reactor in the reactive distillation column from the outflow of a liquid receiver below the lowest reaction zone or from the downcomer of a distillation bottom below the lowest reaction zone. Particularly preferably, the removal of the feed to the side reactor in the reactive distillation column takes place from the downcomer of the collecting tray directly below the lowermost reaction zone. The design of liquid page prints is described in detail in the literature, such. In N. Lieberman, Process Design for Reliable Operation, Gulf Publishing Co., 1988, New York, 1990. Two technical implementations are exemplified by Figures 2a and 2b. It may be advantageous to collect the liquid flowing out of the reaction zone on a chimney tray KB or collecting tray S. In this case, it is advantageous to pass the liquid via a downcomer A into a closure cup V which is positioned somewhat lower than the first distillation bottom below the reaction zone. From this, the side draw stream E is withdrawn via a connection in the column wall. In order to avoid partial evaporation of the boiling mixture in the outlet nozzle, it may be advantageous to design the nozzle for "self-venting flow". By lowering the side discharge cup, the pressure present at the nozzle is increased, so that the risk of partial evaporation is further reduced. For the failure of the side reactor (eg by a control error or pump failure), it is advantageous to ensure a sufficiently large flow from the downcomer A of the chimney tray to the distillation bottom D. In Fig. 2a an exemplary realization is outlined. Very particularly preferably, the removal of the feed E to the side reactor takes place in the reactive distillation column above the feed Z of the reaction mixture obtained from the side reactor into the reactive distillation column. Such a removal is shown in Fig. 2b.
    In the process according to the invention preferably at least 70%, preferably at least 80%, more preferably at least 90%, preferably 95% and very preferably 100% of the internal liquid stream (internal reflux) of the reactive distillation column at the point of removal of the feed to the side reactor as an inlet used for the side reactor.
    The reaction mixture obtained from the side reactor or the last reactor of a series of side reactors is preferably below the reaction zone in the
    Reactive distillation column recycled. Preferably, the reaction mixture above the feed point of the reaction mixture from the pre-reactor is returned to the reactive distillation column. Particularly preferably, the reaction mixture obtained from the side reactor is fed into the downcomer of the collection bottom directly below the reaction zone or into the downcomer of the first distillation bottom directly under the distillation zone or on the second distillation bottom below the distillation zone of the reactive distillation column. If the reaction mixture obtained from the side reactor is fed into a downcomer, from which the feed to the side reactor is taken, then the feed is preferably carried out below the removal point for the feed to the side reactor. Instead of being in a downcomer, the reaction mixture obtained from the side reactor can in particular also be returned to the first bottom adjacent to the underside of the reaction zone.
    side reactor
    The side reactor may have multiple reaction zones or consist of several series-connected partial reactors. As a side reactor or partial reactors conventional fixed bed reactors (tube bundle reactors, adiabatic fixed bed reactors, circulation reactors, etc.) can be used. Preferably, the side reactor is operated in a straight pass. The side reactors have the above-mentioned acidic catalyst. The side reactor may be completely filled with catalyst or have catalyst zones and zones without catalyst. In the process according to the invention, the side reactor is preferably operated with a specific catalyst loading (LHSV) of 0.5 to 20 h'1, preferably of 5 to 15 h'1 and particularly preferably of 10 to 14 h'1.
    The side reactor is preferably operated at a temperature of from 20 to 100 ° C, preferably from 40 to 90 ° C, and more preferably from 50 to 80 ° C. If the side reactor has several reaction zones or if there are several partial reactors, these can be operated at the same or different temperatures. If the reaction zones or partial reactors are operated at different temperatures, it may be advantageous to operate a reaction zone or a partial reactor at a higher temperature than the subsequent reaction zone or the subsequent partial reactor.
    It may be advantageous if the temperature of the feed to the side reactor is adjusted to a temperature of 0 to 40 K, preferably 5 to 30 K and particularly preferably 10 to 20 K smaller than the average temperature in the reaction zone of the reactive distillation column. This can be z. Example, take place in that the inlet to the side reactor before entry into the side reactor by means of one or more cooler (s) is cooled. As a cooler, all known coolers, such. B. air or water cooler can be used. It may be particularly advantageous if, in at least one of the coolers, the feed to the side reactor is cooled against the reaction mixture from the side reactor, which is returned to the reactive distillation column.
    The pressure in the side reactor is preferably set to 0 to 10 bar, preferably 0.5 to 5 bar and particularly preferably 1 to 3 bar above the operating pressure of the reactive distillation column.
    The side reactor is preferably operated so that at least part of the reflux of the reactive distillation column is preferably reacted to or close to the thermodynamic equilibrium.
    In the process according to the invention, the side reactor is preferably operated with a molar ratio of alcohol to isobutene in the feed to the side reactor of 20: 1 to 1: 1, preferably 10: 1 to 1.5: 1 and more preferably 5: 1 to 2: 1 , If the feed into the side reactor at the exit from the reactive distillation column has too low an alcohol content, it can be advantageous if additional alcohol is fed into the feed of the side reactor. It may also be advantageous if additional alcohol is fed into the side reactor itself, in particular directly into the zones of the side reactor having the catalyst.
    With the process according to the invention, in particular as top product of the reactive distillation column, compositions are obtainable which contain at least 25% by mass, preferably 25 to 95% by mass of 1-butene, 0.5 to 7% by mass of alcohol ROH with R = methyl or ethyl, and between 10 and 2000 wppm, preferably between 10 and 600 wppm isobutene. Preferably, the composition contains isobutene in a concentration of less than 2000 wppm based on the 1-butene contained in the composition. Optionally, this composition may contain up to 1% by mass of butadiene.
    The method according to the invention and the apparatus used are explained in more detail below with reference to the FIGS. 1 to 4, without the method being restricted to the embodiment exemplified there.
    Fig. 1 shows schematically an apparatus in which the method according to the invention can be carried out. In Fig. 1 was on the representation of procedural usual streams such. As cooling water streams, circulatory streams or returns, and / or conventional equipment, such. As heat exchangers or separators, partially omitted in favor of a better overview. According to the embodiment of the process according to the invention shown in FIG. 1, a C4-hydrocarbon mixture (for example raffinate I or selectively hydrogenated cracking Cu with alcohol (2a) in the reactor (3) containing an acid catalyst is converted to a mixture containing ATBE, A further amount of alcohol (2b) may be introduced into the catalyst packing (5b) or above the catalyst packing (5b), which also contains an acidic catalyst (9) of the internal liquid stream of the reactive distillation column is reacted, optionally with the addition of further ethanol (2c) in the side reactor, which also contains an acidic catalyst The rectifying mixture (Γ1) leaving the side reactor (tO) is sprayed onto the top bottom of the section ( The top product (7) is an alcohol-containing C4-current, the bez ogen to the 1-butene content less than 2000 wppm isobutene. The bottom product obtained is a mixture which consists of more than 75% of ATBE. Optionally, an alcohol-rich side stream (8) can be withdrawn, which can be returned to the plant, for example in the prereactor (3).
    FIGS. 2a and 2b schematically show implementation possibilities for a side take-off.
    In Fig. 2a an implementation possibility for a single-flow soil is shown. In this case, the liquid flowing out of the reaction zone is collected on a chimney bottom KB (or collecting floor S) equipped with a chimney K. The liquid passes via a downcomer A into the somewhat lowered closure cup V, from which the outlet E for the side reactor is taken off via a connecting piece. The reaction mixture obtained from the side reactor is fed via an inlet Z directly to the distillation bottom D.
    In Fig. 2b shows an implementation possibility for a double-bottomed is shown. In this realization possibility, the liquid flowing out of the reaction zone is collected on a collecting floor S (or chimney bottom KB) equipped with two chimneys K. The liquid is passed into a downcomer A, which has a conclusion with overflow AÜ. Above the conclusion AÜ the deduction E is taken from the side reactor.
    The reaction mixture obtained from the side reactor is passed via an inlet Z below the conclusion of the ATA on the distillation tray D.
    FIG. 3 graphically shows the influence of the side-stream quantity on the residual concentration of isobutene in the distillate, as results from Example 2.
    FIG. 4 graphically shows the influence of the amount of catalyst in the side reactor on the residual concentration of isobutene in the distillate, as it results from Example 2.
    The following examples are intended to illustrate the invention without restricting its scope, which is apparent from the description and the claims.
    Examples
    The following example calculations were carried out with the simulation program ASPEN Plus. In order to generate transparent, reproducible data, only generally accessible material data was used. This makes it easy for the skilled person to understand the calculations. Although the methods used do not have sufficient accuracy for the design of technical systems, the qualitative differences between the two
    Sample calculations will be recorded correctly.
    Example 1 (Comparative Example, Apparatus without Side Reactor)
    For ease of reference, example calculations based on kinetics, equilibrium constant, and vapor-liquid equilibrium literature data were performed in ASPEN Plus version 12.1.
    For kinetics, the approach of Fité et al. 1994 [Fité, C., Iborra, M., Tejero, J., Izquierdo, J.F., and F. Cunill: Kinetics of the Liquid-Phase Synthesis of Ethyl tert-Butyl Ether, Ind. Eng. Chem. Res. 1994, 33, 581-591]. An equilibrium constant published by the same group was used [Vila M., Cunill, F., Izquierdo, JF, Tejero J., and Iborra, M.,: Equilibrium Constants for Ethyl tert-Butyl Ether Liquid-Phase Synthesis , Chem. Eng. Commun., 1993, 124, 223-232]. For the vapor-liquid equilibrium, the modified UNIFAC-Dortmund method is used in ASPEN.
    As a comparative method, a single-stage process was selected consisting of a pre-reactor and a reactive column, as z. B. from EP 1 199 296 for MTBE production is known. The example of the process according to the invention is identical in all parameters with the comparative process, but below the reactive zone of the reactive column, a side reactor is attached.
    The feed used is a stream of 10 t / h of C4 mixture having an isobutene content of 28% by mass and a 1-butene content of 35% by mass in accordance with Table 2, which has an ethanol flow of 2540 kg / h (equivalent to a 10% excess of ethanol to increase the selectivity of the reaction) is mixed. From this starting material mixture, the isobutene is separated by reaction with ethanol to ETBE far enough that in the further workup of this raffinate II stream (ethanol extraction and 2-fold distillation) a 1-butene stream of 2000 kg / h, with a content of 1-butene of 99.7 mass% and an isobutene content of at most 2000 ppm is obtained. Since isobutene is selectively recovered in the 1-butene product during the 1-butene distillation, experience has shown that isobutene accumulates approximately by a factor of 3. To obtain the product specification of 1-butene, a raffinate Π must therefore be used in the ETBE plant
    Concentration of isobutene of a maximum of 600 ppm are generated.
    The educt mixture is first fed to a pre-reactor (for data see Table 3). The pre-reactor is operated as an isothermal reactor at 40 ° C and 10 bar. With a catalyst volume of about 27 m3, the equilibrium conversion of about 91% is then achieved. Thus, the feed concentration to the reactive column is 2.0 mass% isobutene.
    Table 2: Feed and product composition of the pre-reactor
    Table 3: Data of the front and side reactors
    The reactive column has 60 theoretical stages. The feed is introduced after preheating close to its boiling point to the level 50 (counted from above, condenser = stage 1). The reactive zone extends from stage 12 to 34. In the 26th stage, an additional ethanol
    Electricity of 143 kg / h driven. The operating pressure is 10 barabs, the return ratio L / D is 1.
    Sulzer Katapak SP12 [Behrens, M, Olujic, Z., and PJ Jansens: Combining Reaction with Distillation - Hydrodynamic and Mass Transfer Performance of Modular Catalytic Structured Packing, Chemical Engineering Research and Design, 2006, 84 (A5): 381-389]. Accordingly, the package contains about 25% by volume of catalyst. The separation efficiency can be estimated at about 2 theoretical steps per meter of package height. Based on an F factor in the packing of about 1.1 Pa 0.5, a diameter of 1 m was calculated for the reactive zone. The sprinkling density is 17.6 m3 / (m2 * h). These data are summarized in Table 4. The result is a catalyst volume in the reactive zone of about 0.1 m3 per theoretical stage or 2.3 m3 in total.
    Table 4: Catalyst volume in the reactive zone
    In the comparative method, a residual concentration of 910 ppm of isobutene in the distillate is achieved in the reactive column. This results in more than 2700 ppm isobutene in the product 1-butene and the specification is not reached.
    Example 2 (according to the invention)
    The side reactor system is the same reference procedure in all parameters, except here in addition to the liquid collector directly below the reactive zone (step 34) carried out a nearly total side of the withdrawing from the packing liquid. The side draw stream is cooled to 45 ° C and passed through an adiabatically operated fixed bed reactor without recycle. The reactor contains lm3 catalyst (for data see Table 3). The product of the side reactor is introduced after heating to the boiling point immediately below the sampling point at level 35 back into the column.
    In the process according to the invention, a residual concentration of 585 ppm of isobutene in the distillate is achieved in the reactive column with side reactor. Thus, in the product 1-butene at a factor of 3, just under 2000 ppm of isobutene are found, and the specification is achieved (Table 5).
    Table 5: Residual contents of isobutene in the raffinate II and in the product 1-butene
    The distribution of the catalyst volumes on the front and side reactor and the catalytic reaction zone in the packing is shown in Table 6. As expected, both processes require the largest part of the catalyst volume for equilibrium adjustment in the pre-reactor, about 90%. This is due to the isothermal Reaktionsfuhrung at 40 ° C. By a temperature-controlled Reaktionsfuhrung in at least two reactors with falling temperature in the current direction, the catalyst volume of the fixed bed reactors could be reduced to just under 20 m3. As can be seen from the calculation of the process according to the invention, the catalyst volume of the side reactor is about 3.3% of the total volume or 45% of the catalyst volume in the reactive zone of the reactive distillation column.
    Table 6: Catalyst volumes
    FIG. 3 shows the influence of the relative mass flow to the side reactor on the residual concentration of isobutene in the distillate for the calculated example. As the percentage of feed to the side reactor increases, the distillate concentration of isobutene drops monotonously. At 100% relative mass flow of the resulting at the sampling liquid reflux stream is completely removed. The relative mass flow S / L is defined by the ratio of lateral withdrawal amount to return flow at the withdrawal point and is limited to a value of less than or equal to 100% (for reasons of simulation technology). It can be clearly seen that the residual concentration of isobutene in the distillate decreases monotonically with the relative mass flow to the side reactor.
    Analogously, FIG. 4 shows the influence of the size of the side reactor on the residual concentration of isobutene in the distillate. The size of the side reactor is shown here based on the catalyst volume in the reactive packing. It can be clearly seen that even small relative catalyst volumes in the side reactor have a significant influence on the residual concentration of isobutene in the distillate. If more than about 25% of the catalyst volume of the reactive packing is installed in the side reactor, then the residual concentration of isobutene in the distillate approaches a final value. In the calculated example, the residual concentration of isobutene in the distillate no longer appreciably changes with a relative catalyst volume of more than 40%.
    权利要求:
    Claims (41)
    [1]
    A process for the preparation of ATBE and a C4 hydrocarbon stream comprising 1-butene and less than 2000 ppm by mass of isobutene, based on the 1-butene contained in the C4 hydrocarbon stream, from a C4 hydrocarbon mixture (1) containing at least Isobutene and 1-butene contains, by reacting an alcohol ROH with R = alkyl having 1 to 5 carbon atoms (2) with the present in the C4 hydrocarbon mixture isobutene on an acidic catalyst in an apparatus containing at least one pre-reactor (3), a reactive distillation column (5), which has at least one reaction zone (5b) and at least one distillation zone, wherein a distillation zone (5a) is arranged below the reaction zone, and at least one side reactor (10), wherein at least a part of the internal liquid stream of the reactive distillation column as feed ( 9) in the side reactor, the reaction mixture (11) from the side reactor in the reactive distillation column back is passed, the reaction mixture leaving the prereactor (4) is introduced below the reaction zone in the reactive distillation, and being withdrawn as the top product (7) of the reactive distillation column an alcoholic C4 hydrocarbon stream and the bottom product (6) at least 50% by mass ATBE current is characterized in that in the side reactor, a catalyst volume is used, which corresponds to 10 to 160% of the catalyst volume in the reactive distillation column.
    [2]
    2. The method according to claim 1, characterized in that at most 25% by volume of the total catalyst volume present in the apparatus is present in the side reactor.
    [3]
    3. The method according to claim 1 or 2, characterized in that at least 70% of the internal liquid stream of the reactive distillation column is used at the point of removal of the feed to the side reactor as an inlet (9) to the side reactor.
    [4]
    4. The method according to at least one of claims 1 to 3, characterized in that the reaction mixture obtained from a side reactor (11) below the reaction zone (5b) and above the feed point of the reaction mixture from the prereactor (3) is recycled to the reactive distillation column.
    [5]
    5. The method according to at least one of claims 1 to 4, characterized in that the removal of the inlet (9) to the side reactor in the reactive distillation column from the downcomer of the collection bottom below the reaction zone (5b) or from the downcomer of a distillation bottom.
    [6]
    6. The method according to at least one of claims 1 to 5, characterized in that the reaction mixture obtained from a side reactor is fed into the downcomer of the collecting tray below the reaction zone (5b) or in the downcomer of a distillation bottom.
    [7]
    7. The method according to at least one of claims 1 to 5, characterized in that the reaction mixture obtained from a side reactor on the first, the bottom of the reaction zone adjacent bottom, is returned.
    [8]
    8. The method according to at least one of claims 1 to 7, characterized in that the removal of the feed (9) to the side reactor in the reactive distillation above the feed of the reaction mixture obtained from a side reactor (11). Method according to at least one of claims 1 to 8, characterized in that the temperature of the feed to the side reactor is set to a temperature of 0 to 40 K smaller than the average temperature in the reaction zone of the reactive distillation column. Method according to at least one of claims 1 to 9, characterized in that the side reactor is operated with a molar ratio of alcohol to isobutene in the feed to the side reactor of 20 to 1 to 1 to 1. Process according to at least one of Claims 1 to 10, characterized in that the reaction conditions in the at least one pre-reactor are chosen so that the isobutene is reacted with the alcohol at the exit of the reaction mixture from the last prereactor up to or close to equilibrium with ATBE. Method according to at least one of claims 1 to 11, characterized in that the alcohol contained in the top stream of the distillation column is washed out in an extraction step with water. Method according to at least one of claims 1 to 12, characterized in that at least two pre-reactors are present, of which at least one is operated with a return. Method according to at least one of claims 1 to 13, characterized in that in the starting C4 hydrocarbon mixture contained polyunsaturated hydrocarbons, before the reaction of the C4 hydrocarbon mixture with alcohol, are catalytically hydrogenated.
    [9]
    15. The method according to at least one of claims 1 to 14, characterized in that an ion exchange resin is used as the acidic catalyst.
    [10]
    16. The method according to at least one of claims 1 to 15, characterized in that in the reactive distillation column, the acidic catalyst is used in the form of tissue packings which enclose the acidic catalyst.
    [11]
    17. The method according to at least one of claims 1 to 16, characterized in that an alcohol ROH with R = methyl or ethyl is used.
    [12]
    18. The method according to at least one of claims 1 to 17, characterized in that ethanol is used as the alcohol containing from 0.05 to 1% by mass ETBE as denaturant.
    [13]
    19. The method according to at least one of claims 1 to 18, characterized in that the top product obtained C4 hydrocarbon stream is worked up to a 1-butene, which has a content of isobutene of less than 2000 wppm.
    [14]
    20. Composition containing at least 25% by mass of 1-butene, 0.5 to 7% by mass of alcohol ROH with R = methyl or ethyl and between 10 and 2000 wppm of isobutene obtained as top product of the reactive distillation column in a process according to one of claims 1 to 19th
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    引用文献:
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    WO1993019031A1|1992-03-18|1993-09-30|Neste Oy|Process and apparatus for preparing tertiary alkyl ethers|
    WO1994008927A1|1992-10-16|1994-04-28|Chemical Research & Licensing Company|Process for the preparation of etbe|
    JP3433533B2|1994-10-31|2003-08-04|株式会社島津製作所|Gear pump or motor|
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    CN1780803A|2003-11-07|2006-05-31|日本易能有限会社|Method of synthesizing etbe with hydrous ethanol|CN101955418B|2009-12-16|2013-10-16|华东理工大学|Method for preparing ETBE by coupling separation purification|
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    KR102086563B1|2017-01-06|2020-03-09|주식회사 엘지화학|Method for producing methyl tert-butylether|
    FR3077818B1|2018-02-09|2020-02-28|IFP Energies Nouvelles|PROCESS FOR SEPARATING NON-LINEAR OLEFINS FROM AN OLEFINIC FILLER BY REACTIVE DISTILLATION|
    法律状态:
    2019-10-02| MM| Lapsed because of non-payment of the annual fee|Effective date: 20181231 |
    优先权:
    申请号 | 申请日 | 专利标题
    DE102006057856|2006-12-08|
    DE102006057856A|DE102006057856A1|2006-12-08|2006-12-08|Production of alkyl tertiary butyl ether and a hydrocarbon stream containing 1-butene and a small amount of isobutene comprises using a catalyst volume in a side reactor corresponding a catalyst volume in a reactive distillation column|
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