Patent Application
20240417353
This patent application is based on and claims priority pursuant to 35 U.S.C. § 119 (a) to European Patent Application No. 23179203.7, filed on Jun. 14, 2023, in the European Patent Office, the entire disclosure of which is hereby incorporated by reference herein.
The present invention relates to a process for producing a target alcohol by hydrogenation of an ester formed by alkoxycarbonylation of diisobutene. The diisobutene stream used in the alkoxycarbonylation is subjected to a distillation prior to the alkoxycarbonylation in order to enrich 2,4,4-trimethylpent-1-ene in the stream to be alkoxycarbonylated. The alkoxycarbonylation is carried out with an alcohol and carbon monoxide in the presence of a homogeneous catalyst system that comprises at least one metal from group 8 to 10 of the periodic table of the elements or a compound thereof, a phosphorus-containing ligand and an acid.
The production of alcohols in industrial chemistry is achieved largely by hydroformylation to produce an aldehyde, with subsequent hydrogenation of the aldehyde to an alcohol. Although the production of alcohols through hydroformylation with subsequent hydrogenation has been an industrially established and proven process for decades, there is still potential for improvement. One problem with this synthetic route is that pressures and temperatures in the hydroformylation are usually high and hence comparatively high technical demands are made on the plants used. Ultimately, the plants have to be able to withstand the pressures and temperatures.
Diisobutene is a technically relevant product obtained by dimerization of isobutene. Diisobutene consists of the isomers 2,4,4-trimethylpent-1-ene (hereinafter also referred to as TMP1) and 2,4,4-trimethylpent-2-ene (hereinafter also referred to as TMP2) with a mass distribution TMP1: TMP2 of from approx. 78:22 to 81:19 (equilibrium distribution). This mixture can inter alia be converted into higher-value products in carbonylation processes. Particularly in carbonylation processes, such as methoxycarbonylation (reaction product here is methyl 3,5,5-trimethylhexanoate) or hydroformylation (reaction product here is 3,5,5-trimethylhexanal), the internal olefin TMP2 has significantly lower reactivity than the terminal olefin TMP1.As a consequence of inadequate stability of the catalysts or economic factors (for example the size of the reactor), it is often not possible to adjust the reaction conditions or residence times in these carbonylation reactions sufficiently to allow the terminal TMP2 to react to completion.
The object of the present invention was therefore to provide a process for producing a target alcohol that does not have the aforementioned problems. In particular, the alkoxycarbonylation should be carried out with diisobutene streams that have a high proportion of 2,4,4-trimethylpent-1-ene. In addition, an alternative synthesis route for producing a target alcohol should be obtained with which the desired target alcohols can be accessed relatively easily.
This object was achieved by the process of the invention according to the description herein. Preferred embodiments are also specified.
The present invention describes a process for producing a target alcohol, the process comprising at least the following steps:
An advantage of the process is the distillation of the diisobutene stream before the alkoxycarbonylation, since this makes it possible to increase the proportion of 2,4,4-trimethylpent-1-ene in the overhead stream of the at least one distillation column compared to the stream used. Thus, it is possible to achieve proportions of more than 85% by weight, preferably more than 90% by weight, more preferably more than 95% by weight, of 2,4,4-trimethylpent-1-ene in the stream, which is advantageous for the downstream alkoxycarbonylation on account of the higher reactivity of 2,4,4-trimethylpent-1-ene.
Step a of the process of the invention relates to the distillation of a diisobutene stream comprising 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene. The diisobutene stream is separated in this step into at least one overhead stream enriched with 2,4,4-trimethylpent-1-ene compared to the diisobutene stream used that preferably contains more than 85% by weight of 2,4,4-trimethylpent-1-ene, further preferably more than 90% by weight of 2,4,4-trimethylpent-1-ene, particularly preferably more than 95% by weight of 2,4,4-trimethylpent-1-ene, and a residual stream depleted in 2,4,4-trimethylpent-1-ene compared to the diisobutene stream used.
The diisobutene stream used in step a comprises 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene. Such streams may be diisobutene streams which can be produced by dimerization from isobutene or isobutene-containing hydrocarbon mixtures, for example those disclosed in EP 1 360 160 B1. In addition, the streams to be used here may occur as unreacted residual streams in carbonylation processes, for example in a methoxycarbonylation or in a hydroformylation. It has already been mentioned that the input stream can be supplied to the isomerization and/or to the distillation. The designation of steps a. and b. in the present description therefore does not constitute a prioritization or chronology. It will be apparent that both steps must be carried out and that the individual steps are fed from the respective other step. Which step is therefore considered to be carried out first is therefore of secondary importance.
The distillation in step a of the process of the invention is carried out in at least one distillation column. Distillation columns are generally known to those skilled in the art. The distillation unit of the present invention comprises preferably at least one distillation column, more preferably at least two distillation columns. The following description of features of the distillation column also applies whenever there is more than one distillation column in the distillation unit.
The at least one distillation column preferably has internals to cope with the separation task. Appropriate internals are familiar to those skilled in the art. Particularly suitable here are random or structured packings, such as those known to those skilled in the art under trade names such as MellaPak®, MellapakPlus®, Flexipac®, etc. In a preferred embodiment of the present invention, the at least one distillation column comprises at least 50 theoretical plates, preferably at least 70 theoretical plates, more preferably at least 90 theoretical plates. If two or more distillation columns are present, the distillation columns may have an identical or different number of theoretical plates.
The operating parameters of the at least one distillation column are oriented to the separation task and the design of the distillation column. In the present process, it is preferable that the at least one distillation column is operated at reduced pressure, particularly preferably at a pressure of 0.2 to 0.9 bar. In the context of the present invention, reduced pressure is present whenever operations are carried out below the ambient pressure of approx. 1 bar, i.e. the atmospheric pressure present at the respective site. If there are two or more distillation columns, the distillation columns may be operated at the same or different pressure.
It is additionally preferable that the temperature in the bottoms of the at least one distillation column is in the range from 50 to 100° C. If there are two or more distillation columns, the distillation columns may be operated at the same or different temperature.
A further parameter for the design of distillation columns is the reflux ratio. The reflux ratio means the ratio of reflux (recyclate) to distillate (withdrawn condensate, in this case therefore the overhead stream K). The reflux is thus a condensed portion of the overhead stream that is returned to the distillation column. It is preferable in accordance with the invention that the reflux ratio of the at least one distillation column in the distillation in step a. is in the range from 5 to 15.
In the at least one distillation column, the distillation according to the invention in step a gives rise to an overhead stream that preferably contains at least 85% by weight of 2,4,4-trimethylpent-1-ene, further preferably at least 90% by weight of 2,4,4-trimethylpent-1-ene, particularly preferably at least 95% by weight of 2,4,4-trimethylpent-1-ene. In addition, the distillation affords a residual stream depleted in 2,4,4-trimethylpent-1-ene compared to the diisobutene stream used.
The residual stream additionally obtained in the distillation in step a can be withdrawn as a bottoms stream or as a side stream of the at least one distillation column. If there are two distillation columns, the residual stream is obtained in the final distillation column, irrespective of whether it is withdrawn as a side stream or bottoms stream. If the residual stream can be withdrawn as a side stream, this may in principle be arranged at the same height or below the inlet for the employed diisobutene stream. It will be apparent here that the inflow and outflow must be arranged an adequate distance apart. The residual stream, irrespective of whether it is withdrawn as a side stream or bottoms stream, preferably contains 80% to 92% by weight of 2,4,4-trimethylpent-2-ene.
If the residual stream in the distillation in step a is withdrawn as a bottoms stream of the at least one distillation column, a purge stream comprising high-boilers formed during the isomerization, for example dimers or oligomers of 2,4,4-trimethylpent-2-ene and/or 2,4,4-trimethylpent-1-ene, is preferably withdrawn from the residual stream. The removed high boilers can be incinerated to generate energy or can undergo hydrogenation to produce valuable alkanes.
If the residual stream in the distillation in step a is withdrawn as a side stream of the at least one distillation column, a stream comprising high-boilers formed during the isomerization, for example dimers or oligomers of 2,4,4-trimethylpent-2-ene and/or 2,4,4-trimethylpent-1-ene, is preferably obtained in the bottoms of the at least one distillation column. The removed high boilers can be incinerated to generate energy or can undergo hydrogenation to produce valuable alkanes.
Distillation columns are, as is known, heated in the bottoms to cope with the separation task. It is possible here to supply energy via heating steam, which is often available at chemical production sites. This is done by passing a portion of the bottoms through a heat exchanger (reboiler), heating it there and then returning it to the bottoms of the distillation column. To reduce energy requirements and/or CO2 emissions, the distillation in step a may be designed with thermal integration. Thermal integration means that energy produced or present within the process is used elsewhere. In the present case, the (released) condensation energy arising at the top of at least one distillation column of the distillation unit is particularly suitable. In a preferred embodiment of the present invention, thermal integration is carried out in step a. in such a way that at least part of the condensation energy at the top is used for heating the bottoms.
Another option for thermal integration is what is known as vapour compression, in which at least a portion of the vapour (overhead stream) is compressed, i.e. pressurized to a higher level and optionally heated. The vapour thus compressed and optionally heated is supplied to the heat exchanger (reboiler) in order to heat the bottoms. This makes use of the heat of condensation of the vapour.
A further option is the use of a heat pump. Heat pumps are operated using a working medium such as n-butane or water. The condensation energy is in this case thus first transferred in a heat exchanger to a working medium and from there, in a further suitable heat exchanger, to the distillation bottoms. The working medium is here usually conveyed via a compressor to the heat exchanger in the bottoms. It is in principle also possible to employ two-stage or multistage heat pumps that have more than one compressor stage. In the case of a two-stage heat pump, there is not just one working medium, but two working media, wherein an exchange of energy between the first and the second working medium also takes place in a heat exchanger. In the case of multistage configurations, correspondingly more working media are present.
The overhead stream obtained from the distillation in step a is supplied to the subsequent step b of alkoxycarbonylation and reacted there. In a preferred embodiment, the residual stream also obtained from the distillation in step a is supplied to an isomerization in which 2,4,4-trimethylpent-2-ene is at least partially isomerized to 2,4,4-trimethylpent-1-ene using a heterogeneous catalyst based on a zeolite or on an ion-exchange resin. The isomerization can thus be carried out with the residual stream depleted in 2,4,4-trimethylpent-1-ene compared to the diisobutene stream used in step a and withdrawn from the distillation in step a.
The performance of an isomerization prior to the distillation in step a has the advantage that a heterogeneous catalyst system is used. Such a catalyst system does not have to be removed from the isomerization stream and remains in the reaction vessel/reactor. The catalyst systems based on a zeolite or on an ion-exchange resin that are preferred in accordance with the invention additionally permit high conversions to be achieved.
The isomerization can in principle be carried out in any suitable reactor. It is possible for the isomerization to take place in a single reactor or in two or more reactors connected in parallel or in series. Execution in batches or in continuous operation is also possible. The isomerization is preferably carried out in one or more continuously operated reactors that are customarily employed in solid/liquid contact reactions. When using continuously operated flow reactors, a fixed bed is usually, but not always, employed. When a fixed-bed flow reactor is used, the liquid can flow in an upward or downward direction. In most cases, downward flow of the liquid is preferable. In addition, it is possible to operate the reactor with product recycling or in straight pass. A concept different to that of fixed-bed reactors is for example reactors in which the ion exchanger or zeolite is suspended in a liquid phase.
The reactors used for the isomerization may be tubular reactors or tube bundle reactors, especially ones having internal tube diameters of 10 to 60 mm. The length-to-diameter ratio of the catalyst bed may be varied here, either by the geometric dimensions of the reactor or by its filling level. At the same amount of contact and load (LHSV), it is thus possible to achieve different superficial velocities and to selectively influence the heat transfer to the cooling medium.
The cooling of the tubes of the reactor, whether it be a tubular reactor or a tube bundle reactor, can be effected via a cooling medium (for example cooling water or a heat-absorbing process fluid for thermal integration) via the shell space of the reactor or a heat exchanger in an external recycling system. Especially when using liquid heating media, the shell side is constructed such that the temperature gradient in contact with all tubes is as homogeneous as possible. The technical measures necessary for this are known to those skilled in the art and are described in the literature (installation of baffle plates, disc-on-donut construction, infeed/outfeed of heat-transfer medium at various points in the reactor, etc.). Preferably, the reaction medium and heat transfer medium are respectively conveyed through the reactor tubes and reactor jacket in cocurrent flow, more preferably from top to bottom. A preferred embodiment is described for example in DE 10 2006 040 433 A1.
The isomerization of 2,4,4-trimethylpent-2-ene to 2,4,4-trimethylpent-1-ene takes place exothermically, i.e. it proceeds with the release of energy, which results in warming of the reaction mixture. In order to limit the temperature rise, it is possible to dilute the input stream, for example by recycling the product.
The reactor(s) used in the isomerization may be operated adiabatically, polytropically or practically isothermally. Practically isothermally means that the temperature is at no point in the reactor more than 10° C. higher than the temperature at the reactor entrance. In the case of adiabatic operation of the reactors, it is usually advantageous to arrange a plurality of reactors in series and to provide cooling between the reactors. Reactors that are suitable for polytropic or practically isothermal operation are for example the tube bundle reactors already mentioned and also stirred-tank reactors and loop reactors.
The process can be executed at rather mild temperatures. The isomerization is preferably carried out at a temperature of from 25 to 90° C., preferably 30 to 80° C., more preferably 35 to 70° C. In addition, the isomerization of the invention can be carried out at a pressure equal to or greater than the vapour pressure of the input stream mixture and/or of the reaction mixture at the respective reaction temperature, preferably at a pressure of more than 0 bar but less than 40 bar. Further preferably, the isomerization is carried out in the liquid phase. It should be clear that in this case the pressure and temperature must be chosen such that the input stream is present, or may be present, in the liquid phase.
In the reactor(s), it is also possible to use various catalysts based on a zeolite or on an ion-exchange resin for the isomerization. For example, a mixture of ion-exchange resins of different reactivity may be used. It is likewise possible for a reactor to contain catalysts of different activities that are arranged for example in layers. If more than one reactor is used, the individual reactors may be filled with the same or different catalyst(s) based on a zeolite or on an ion-exchange resin.
As a heterogeneous catalyst, a catalyst based on a zeolite or on an ion-exchange resin may be used for the isomerization. Suitable zeolites and ion-exchange resins are widely available on the market. It has been found to be preferable when a zeolite-based catalyst has an Si:Al ratio in the range from 40:1 to 200:1. In the case of catalysts based on an ion-exchange resin, preferably the styrene-divinylbenzene type as the H-form and in a partially neutralized form has in addition been found to be highly suitable.
When the catalyst is a zeolite, some zeolites have proven to be particularly advantageous, for example beta-and gamma-zeolites. The zeolite is therefore preferably selected from the group consisting of Z-beta-H-25, Z-beta-H-38, Z-beta-H-360, Z-Mor-H-20, Z-Y-H-60, Z-Y-H-80, Z-beta-H, Z-CFG-1, Z-beta-ammonium-38, Z-CBV 760 CY (1.6), Z-CBV 780 CY (1.6), CP 814E CY (1.6), CBV 500 CY (1.6), H-CZB-150 and mixtures thereof.
Ion-exchange resins that may be used include for example ion-exchange resins produced by the sulfonation of phenol/aldehyde condensates or by the sulfonation of copolymers of aromatic vinyl compounds. Examples of aromatic vinyl compounds for the production of the copolymers are: styrene, vinyltoluene, vinylnaphthalene, vinylethylbenzene, methylstyrene, vinylchlorobenzene, vinylxylene and divinylbenzene. Particular preference is given to using for the isomerization ion-exchange resins produced by the sulfonation of copolymers formed by the reaction of styrene with divinylbenzene. The ion-exchange resins may be produced in gel-like, macroporous or sponge-like form. The properties of these resins, in particular specific surface area, porosity, stability, swelling or shrinkage and exchange capacity may, as is known, be varied via the production process.
Ion-exchange resins of the preferred styrene-divinylbenzene type are sold inter alia under the following trade names: CT 151 and CT275 from Purolite, Amberlyst® 15, Amberlyst® 35, Amberlite® IR-120, Amberlite® 200 from Rohm&Haas, Dowex M-31 from Dow, Lewatit® K 2621, Lewatit® K 2431 from Lanxess.
The pore volume of the ion-exchange resins employable as catalysts, in particular those of the preferred styrene-divinylbenzene type, is preferably 0.3 to 0.9 ml/g, more preferably 0.5 to 0.9 ml/g. The pore volume can be determined for example by adsorptive techniques. The particle size of the ion-exchange resins is preferably from 0.3 mm to 1.5 mm, more preferably 0.5 mm to 1.0 mm. A narrower or broader particle size distribution may be chosen. It is thus possible for example to use ion-exchange resins having a very uniform particle size (monodisperse resins).
The ion-exchange resins employable as catalysts for the isomerization may be present as partially neutralized ion-exchange resins. For this purpose, the ion-exchange resin may be treated with acids or bases, as described in EP 1 360 160 B1.
The optional isomerization affords an isomerization stream in which the proportion of 2,4,4-trimethylpent-2-ene is lower and the proportion of the 2,4,4-trimethylpent-1-ene in the isomerization stream higher than in the residual stream used. The isomerization stream may be supplied at least in part to the distillation in step a, preferably the entire isomerization stream is supplied to the distillation in step a.
The process may also include the supply of an input stream with which the process is fed. The input stream with which the process is fed and which comprises 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene can be conducted to the distillation unit in step a in addition to the isomerization stream and/or to the isomerization in addition to the residual stream.
The flexible supply of the input stream is advantageous here. This allows the process to be adapted to the respective conditions, for example according to the input stream that is employed. If the proportions of 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene in the input stream are close to the equilibrium distribution, the stream can be supplied directly to the distillation. If the proportions of 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene in the input stream differ from the equilibrium distribution, the stream could initially be supplied to the isomerization. Incorporation into existing production plants is thus possible in a straightforward manner according to the composition of the input stream.
The overhead stream obtained from the distillation in step a is, as mentioned, supplied to the subsequent step b of alkoxycarbonylation and reacted there.
The diisobutenes, i.e. 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene, are in step b reacted with carbon monoxide (CO) and an alcohol to form an ester. This increases the number of carbon atoms in the ester, compared to the diisobutene used, by 1 carbon atom from the carbon monoxide and by the number of carbon atoms in the alcohol used. The diisobutenes (8 carbon atoms) accordingly give rise to an ester having 9 carbon atoms in the acid part of the ester, plus the carbon atoms of the alcohol in the alcohol part of the ester.
The carbon monoxide may in step b be provided directly in the form of a feedstock mixture or through the addition of a carbon monoxide-containing gas selected from synthesis gas, water gas, generator gas and other carbon monoxide-containing gases. It is also possible to provide the carbon monoxide by first separating the carbon monoxide-containing gas into its components in a manner known to those skilled in the art and passing the carbon monoxide into the reaction zone. The carbon monoxide may still contain a certain proportion of hydrogen or other gases, because complete separation is almost impossible.
The alcohol used in the reaction in step b is preferably a monohydric alcohol having 1 to 4 carbon atoms, especially methanol, ethanol, propanol or butanol. This then affords the esters with the employed alcohol of the corresponding chain length. For example, when methanol is used, the methyl ester is formed.
The alcohol used in step b, preferably the monohydric alcohol having 1 to 4 carbon atoms, especially methanol, ethanol, propanol or butanol, is used in a molar ratio to the total amount of all diisobutenes, alcohol: diisobutenes, within a range of from 10:1 to 1:1, preferably 8:1 to 1.5:1, more preferably 7:1 to 2:1. The alcohol is thus added in an at least identical molar amount based on diisobutenes used, but preferably in a molar excess.
The reaction in step b according to the invention is carried out in the presence of a homogeneous catalyst system that comprises at least one metal from group 8 to 10 of the periodic table of the elements (PTE) or a compound thereof, a phosphorus-containing ligand and an acid as co-catalyst. The content of the metal of group 8 to 10 of the periodic table of the elements, in particular of palladium, in the alkoxycarbonylation reaction solution in step b is preferably 100 to 500 ppm, further preferably 150 to 450 ppm, particularly preferably 180 to 350 ppm.
The metal from group 8 to 10 of the PTE is preferably palladium. The palladium is preferably used in the form of a precursor compound as a palladium compound coordinated by the phosphorus-containing ligand. Examples of palladium compounds that may be used as precursor compounds are palladium chloride [PdCl2], palladium(II) acetylacetonate [Pd(acac)2], palladium(II)acetate [Pd(OAc)2], dichloro(1,5-cyclooctadiene)palladium(II) [Pd(cod)2Cl2], bis(dibenzylideneacetone)palladium(0) [Pd(dba)2], tris(dibenzylideneacetone)dipalladium(0) [Pd2(dba)3] bis(acetonitrile)dichloropalladium(II) [Pd(CH3CN)2Cl2], palladium(cinnamyl)dichloride [Pd(cinnamyl)Cl2]. Preference is given to using the compounds [Pd(acac)2] or [Pd(OAc)2]. The concentration of palladium metal in step b is preferably between 0.01 and 0.6 mol %, preferably between 0.03 and 0.3 mol %, more preferably between 0.04 and 0.2 mol %, based on the molar amount of the hydrocarbon used.
Suitable phosphorus-containing ligands of the catalyst system according to the invention preferably have a bidentate structure. Preferred phosphorus-containing ligands for the catalyst system according to the invention are benzene-based diphosphine compounds, as disclosed for example in EP 3 121 184 A2. The ligands may be combined with the palladium in a preliminary reaction so that the palladium-ligand complex is fed into the reaction zone, or added to the reaction in situ and combined with the palladium there. The molar ratio of ligand to metal for the reaction described in step b may be 1:1 to 10:1, preferably 2:1 to 6:1, more preferably 3:1 to 5:1.
The homogeneous catalyst system further comprises an acid, in particular a Brønsted or a Lewis acid. Lewis acids used are preferably aluminium triflate, aluminium chloride, aluminium hydride, trimethylaluminium, tris (pentafluorophenyl) borane, boron trifluoride, boron trichloride or mixtures thereof. Of the Lewis acids mentioned, preference is given to using aluminium triflate. The Lewis acid is preferably added in a molar ratio of Lewis acid: ligand of 1:1 to 20:1, preferably 2:1 to 15:1, particularly preferably 5:1 to 10:1.
Suitable Brønsted acids preferably have an acid strength of pKa≤5, more preferably an acid strength of pKa ≤ 3. The stated acid strength pKa relates to the pKa determined under standard conditions (25° C., 1.01325 bar). For a polyprotic acid, the acid strength pKa relates in the context of this invention to the pKa of the first protolysis step. The Brønsted acid is preferably added in a molar ratio of Brønsted acid: ligand of 1:1 to 15:1, preferably 2:1 to 10:1, more preferably 3:1 to 5:1.
The Brønsted acid used may in particular be perchloric acid, sulfuric acid, phosphoric acid, methylphosphonic acid or sulfonic acids. Suitable sulfonic acids are for example methanesulfonic acid, trifluoromethanesulfonic acid, tert-butanesulfonic acid, p-toluenesulfonic acid (PTSA), 2-hydroxypropane-2-sulfonic acid, 2,4,6-trimethylbenzenesulfonic acid and dodecyl sulfonic acid. Particularly preferred acids are sulfuric acid, methanesulfonic acid, trifluoromethanesulfonic acid and p-toluenesulfonic acid. The acid is preferably sulfuric acid.
The alkoxycarbonylation in step b is preferably carried out at a temperature in the range from 60 to 120° C., further preferably in the range from 65 to 110° C., particularly preferably in the range from 70 to 100° C. The alkoxycarbonylation in step b is further preferably carried out at a carbon monoxide pressure of from 10 to 35 bar, preferably 12.5 to 30 bar, more preferably 15 to 25 bar.
The alkoxycarbonylation in step b takes place in a suitable reaction zone. The reaction zone for the reaction comprises at least one reactor, but may also consist of two or more reactors arranged in parallel or in series. The at least one reactor may in particular be selected from the group consisting of a stirred-tank reactor, a loop reactor, a jet-loop reactor, a bubble-column reactor or combinations thereof. If more than one reactor is used, the reactors may be identical or different from one another.
The alkoxycarbonylation in step b described above affords a liquid product mixture that comprises at least the ester formed by the alkoxycarbonylation, the homogeneous catalyst system, the unreacted diisobutenes 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene, and unreacted alcohol. In addition, the liquid product mixture may contain low-boiling by-products such as dimethyl ether or formic acid and/or high-boiling components such as ligand decomposition products.
The product mixture thus obtained is supplied to the following step c in order to remove the homogeneous catalyst system from the liquid product mixture. Before the product mixture is fed in, low-boiling components, for example a portion of the unreacted alcohol and/or of the low-boiling by-products, can be removed. The alcohol can be recycled to step b, i.e. to the alkoxycarbonylation. The low-boiling by-products can be discharged from the process in the sense of a purge, so that they do not accumulate in the process. The removal of the homogeneous catalyst system affords a crude product mixture that comprises at least the esters formed by the alkoxycarbonylation, unreacted diisobutenes and at least a portion of the unreacted alcohols.
The removal of the homogeneous catalyst system to obtain the crude product mixture in step c may be effected with the aid of various separation processes, for example by thermal separation and/or by membrane separation. Suitable processes are familiar to those skilled in the art. The removal of the homogeneous catalyst system in step c is in the context of the present invention preferably effected by membrane separation. As is known, a membrane separation gives rise to a retentate and a permeate. The homogeneous catalyst system will accumulate in the retentate. The permeate in that case constitutes the crude product mixture mentioned previously and is conducted to the subsequent step d of distillative processing.
It is preferable in accordance with the invention that the retentate is recycled to the alkoxycarbonylation in step b/to the reaction zone where the alkoxycarbonylation is carried out. This allows the catalyst system to be reused. In the preferably continuous execution of the described process, this gives rise to a catalyst cycle in which at most only minor process-related catalyst losses need to be compensated. Where the diisobutene stream, the alcohol and the homogeneous catalyst system are, in accordance with the preferred embodiment, mixed more particularly in a suitable mixing vessel prior to the alkoxycarbonylation in step b, the retentate is supplied to the mixing vessel. In the recycling of the retentate, a purge stream that may contain inert alkanes, low-boiling by-products (for example ethers), possible decomposition products of the catalyst system or other impurities, for example traces of water or nitrogen, may additionally be withdrawn to avoid accumulation in the reaction zone(s).
Any suitable membrane material may be used for the membrane separation. Preference is given to using an OSN (organic solvent nanofiltration) membrane material in the membrane separation in step c of the process of the invention. Such a membrane material preferably consists at least of a separation-active layer (also: active separation layer) and a substructure on which the separation-active layer is present. The membrane material of the invention preferably consists at least of a separation-active layer and a substructure.
The membrane material, composed at least of separation-active layer and a substructure, should be acid-stable so that the membrane material is not damaged by the acid present in the liquid product mixture. The term “acid stable” means in the context of the present invention that the membrane material is stable for at least 300 h in the presence of the acid of the catalyst system and the separation performance is maintained.
The substructure preferably has a porous structure that is permeable to the permeate that has passed through the separation-active layer. The substructure has a stabilizing function and serves as a support for the separation-active layer. The substructure may in principle consist of any suitable porous material. Suitable materials are familiar to those skilled in the art. A prerequisite, however, is that the material is stable to acids and bases. The substructure may consist of the same material as the separation-active layer. Preferred materials for the substructure are plastics such as polypropylene or a PAEK polymer (PAEK =polyaryletherketone). The PAEK polymer is in particular PEEK (polyether ether ketone), more preferably a PEEK polymer having a degree of sulfonation of less than 20%, more preferably having a degree of sulfonation of less than 10%.
The separation-active layer according to the invention is preferably composed of a PAEK (polyaryletherketone) polymer. PAEK has the particular feature that, within the repeat unit, aryl groups are linked alternately via an ether functionality and a ketone functionality. A separation-active layer that is preferred according to the invention is composed of PEEK (polyether ether ketone). As the separation-active layer, particular preference is given to using PEEK polymers having a degree of sulfonation of less than 20%, particularly preferably having a degree of sulfonation of less than 10%. The corresponding PEEK polymers and the production thereof are described in WO 2015/110843 A1.
The membrane separation in step c is preferably carried out at a temperature in the range from 25° C. to 100° C., further preferably in the range from 30° C. to 80° C. and particularly preferably in the range from 40° C. to 70° C. To bring the product mixture to the prevailing temperature preferred for the membrane separation, the product mixture may be cooled. In addition to active cooling using a coolant, cooling may also be achieved via a heat exchanger, whereby another stream is heated within the process of the invention.
The transmembrane pressure (TMP) in the membrane separation in step c is preferably in the range from 10 to 60 bar, further preferably in the range from 15 to 55 bar, particularly preferably in the range from 20 to 50 bar. The permeate-side pressure here may be above atmospheric pressure and preferably up to 15 bar, preferably 3 to 7 bar. The difference between the TMP and the permeate-side pressure gives the retentate-side pressure. In a preferred embodiment, care should be taken, in the case of the pressure ratios and the permeate-side pressure in particular, to ensure that the pressure is set in accordance with the hydrocarbon used, the alcohol used and the temperature in the system, such that evaporation after passage through the membrane is avoided. Evaporation could lead to unstable operation.
In the subsequent step d, the distillative processing of the crude product mixture/permeate from the membrane separation is carried out in at least one distillation column to remove the unreacted alcohols and the unreacted diisobutenes. This affords an ester product comprising the esters formed. The ester product comprises the esters formed from the diisobutene, i.e. the alkyl 3,5,5-trimethylhexanoate esters, in which the alkyl radical contains 1 to 4 carbon atoms depending on the alcohol used.
In the distillative processing of the crude product mixture/permeate in step d, the unreacted alcohol and the unreacted diisobutenes are obtained at the top of the at least one distillation column. The ester product thus accumulates in the bottoms of the at least one distillation column. The overhead stream comprising the unreacted alcohol removed in the at least one distillation column and the unreacted diisobutenes is recycled to the alkoxycarbonylation in step b. If the components used undergo mixing prior to the alkoxycarbonylation, the overhead stream is by definition supplied to the mixing. This allows continuous operation of the process of the invention at the highest possible yield. A purge can be withdrawn from the recycled overhead stream in order to discharge low-boiling by-products from the process.
The distillative processing to remove the unreacted alcohols and the unreacted diisobutenes in step d can be carried out in a distillation column. It would be conceivable for the distillative processing for the removal of the unreacted alcohols and the unreacted diisobutenes in step d to be carried out in more than one distillation column. This would however mean a significant higher outlay on equipment. It is therefore preferable that the distillative processing in step d is carried out in a single distillation column.
The pressure in the distillation column in the distillative processing in step d is preferably in the range from 0.3 to 2 bar, further preferably in the range from 0.4 to 1 bar, particularly preferably in the range 0.5 to 0.7 bar. The temperature in the bottoms of the distillation column in the distillative processing in step d is preferably in the range from 80°° C. to 160° C. The temperature at the top of the distillation column in the distillative processing in step d is preferably in the range from 30 to 80° C. In addition, it is preferable that the reflux ratio in the distillation column is between 1 and 2. The distillation column for the removal in step d comprises preferably 10 to 30 theoretical plates. The distillation column may contain high-performance structured packings. Suitable high-performance structured packings are known to those skilled in the art.
The ester produced in step b and removed and optionally purified in step d then undergoes a hydrogenation in step e. The ester group is cleaved by the hydrogen used in this process, resulting in the formation of the target alcohol and allowing the alcohol that is bound in the ester formation in step b to be recovered. The hydrogenation accordingly gives rise to an alcohol mixture that comprises at least the target alcohol, the alcohol that is eliminated and unreacted esters.
The alcohol recovered during the hydrogenation in step e can in a subsequent process step be removed from the resulting alcohol mixture and recycled to the first reaction zone.
Hydrogenation is a well-known reaction. The typical hydrogenation conditions are known to those skilled in the art and are disclosed for example in EP 1 042 260 A1. The hydrogenation according to the invention in step e is preferably carried out at a pressure of 10 to 300 bar, further preferably at a pressure of 100 to 280 bar and particularly preferably at a pressure of 150 to 270 bar. The hydrogenation temperature is preferably between 100° C. and 300° C., preferably between 150° C. and 280° C. and particularly preferably between 180° C. and 220° C. In the hydrogenation, an offgas can additionally be withdrawn to remove low-boiling components, for example impurities such as nitrogen or low-boiling by-products.
The hydrogenation in step e takes place in the presence of a heterogeneous catalyst system. Typical catalyst systems are known for example from EP 1 338 557 A1. The heterogeneous catalyst system preferably comprises a metal from the group consisting of copper, rhodium, ruthenium, rhenium or compounds of these metals. In addition, catalyst systems based on copper-chromium oxide are also suitable. Particularly preferred catalyst systems contain copper and/or zinc as the active component, which are applied to a support material. Porous silica and/or alumina are suitable as the support material.
The hydrogen required for the hydrogenation may be supplied directly as feedstock. It is also possible to supply the hydrogen by first separating a hydrogen-containing gas into its components in a manner known to those skilled in the art and conveying the hydrogen to the hydrogenation zone. The hydrogen may still contain a certain proportion of carbon monoxide or other gases, because complete separation is almost impossible.
In the subsequent step f to remove the target alcohol formed in step e from the rest of the alcohol mixture, the alcohol mixture from the hydrogenation in step e is subjected to at least one separation process step selected from the group consisting of a thermal separation, for example distillation, an extraction, a crystallization or a further membrane separation. The separation process is preferably a distillation. The appropriate process conditions are known to those skilled in the art. It is also possible to perform a multistage distillation.
In addition, the alcohol removed in the at least one separation process step that was used in the alkoxycarbonylation can be recycled to the reaction zone. During recycling, a purge stream may be withdrawn in order to discharge from the process hydrogenation by-products such as inert alkanes, aldehydes, acetals, ethers, ketones or carbons.
The present invention is elucidated more particularly with reference to FIG. 1 and FIG. 2. The drawings show specific embodiments and are not intended to limit the subject matter of the invention.
FIG. 1 shows an embodiment in which the diisobutene stream (1) is supplied to the distillation column (2) in step a. The overhead stream (3) is enriched with 2,4,4-trimethylpent-1-ene compared to the diisobutene stream used. The distillation also gives rise to the residual stream (4) that is depleted in 2,4,4-trimethylpent-1-ene compared to the diisobutene stream used. The overhead stream (3) is subsequently supplied to the alkoxycarbonylation (5) in step b, resulting in a product mixture (6). This product mixture (6) undergoes catalyst removal (7) in step c to afford the crude product mixture (8), which is subjected to a distillation (9) in step d. In this distillation, the product (11) and the unreacted alcohols and the unreacted 2,4,4-trimethylpent-2-ene and 2,4,4-trimethylpent-1-ene are obtained as a stream (10) that is recycled to the alkoxycarbonylation (5). The product, i.e. the alkyl 3,5,5-trimethylhexanoate, is then conveyed to the hydrogenation (14) and converted there into the target alcohol, which is removed as a stream (16) in the downstream distillative separation (15).
FIG. 2 relates to a further embodiment in which an isomerization (12) is additionally included. The employed diisobutene stream (13) can here be supplied to the isomerization and/or to the distillation (2). This is indicated by the dashed lines. The isomerized stream (1) then becomes the inflow to the distillation (2). The isomerization is also fed with the residual stream (4) from the distillation, in order there to convert 2,4,4-trimethylpent-2-ene at least partially into 2,4,4-trimethylpent-1-ene.
The invention is described hereinbelow by reference to examples. These are provided for elucidation purposes and do not limit the subject matter of the invention. Claims
| Number | Date | Country | Kind |
|---|---|---|---|
| 23179203.7 | Jun 2023 | EP | regional |
