Optimized separation technique for work-up of homogeneously catalysed hydroformylation mixtures

Abstract

The invention relates to a method for producing alcohols by homogeneously catalyzed hydroformylation of olefins to aldehydes and subsequent hydration of the aldehydes. The invention further relates to a system for carrying out the method. The main focus is on the separation technique for work-up of the hydroformylation mixture. The problem addressed by the invention is that specifying a work-up method for hydroformylation mixtures that utilizes the specific advantages of known separation technologies but at the same time largely avoids the specific disadvantages of said separation technologies. The most important objective is to create a catalyst separation system that is as complete and at the same time conservative as possible and that operates in a technically reliable manner and entails low investment and operating costs. The method should be unrestrictedly suitable for processing the reaction output from oxo systems in world scale format. The problem is solved by combining membrane separation units and a thermal separation unit, the thermal separation unit being operated in such a manner that 80% to 98% of the mass introduced with the product stream into the thermal separation unit exits the thermal separation unit again as a head product.

Claims

1. A process for producing an alcohol, the process comprising: a) providing at least one olefin, syngas and a catalyst system and optionally a solvent; b) hydroformylating the olefin in the presence of the syngas and of the catalyst system in at least one hydroformylation reactor in a homogeneously catalysed hydroformulation to form at least one aldehyde and at least one high boiler; c) withdrawing a liquid hydroformylation effluent comprising the aldehyde, the olefin, dissolved syngas, the catalyst system and the high boiler from the hydroformylation reactor; d) optionally devolatilizing the liquid hydroformylation effluent; e) separating the liquid hydroformylation effluent in a first membrane separation unit into a product stream and a reactor return stream, wherein the catalyst system partitions into the reactor return stream; f) returning the reactor return stream into the hydroformylation reactor; g) optionally devolatilizing the product stream; h) separating the product stream in a thermal separation unit into a gaseous head product comprising a first part of the aldehyde and the olefin and a liquid bottom product comprising a second part of the aldehyde, the high boiler and a catalyst complex; and i) separating the liquid bottom product in a second membrane separation unit into a permeate and a retentate, wherein the catalyst system partitions into the retentate; and wherein an improvement of the process comprises: (j) operating the thermal separation unit such that 80% to 98% of mass introduced into the thermal separation unit with the product stream re-emerges from the thermal separation unit as head product; (k) subjecting at least some of the head product of the thermal separation unit and the permeate of the second membrane separation unit to conjoint or separate hydrogenation; and (l) the second membrane separation unit takes a form of a two-stage enriching cascade.

2. The process according to claim 1, wherein the retentate of the second membrane separation unit is fed to the first membrane separation unit in admixture with the liquid hydroformylation effluent withdrawn from the hydroformylation reactor.

3. The process according to claim 1, wherein the permeate of the second membrane separation unit passes through an adsorber before the hydrogenation.

4. The process according to claim 1, wherein the head product of the thermal separation unit passes through an adsorber before the hydrogenation.

5. The process according to claim 3, wherein the head product of the thermal separation unit and the permeate of the second membrane separation unit pass through the same adsorber before the hydrogenation.

6. The process according to claim 5, wherein the head product of the thermal separation unit and the permeate of the second membrane separation unit are subject to the conjoint hydrogenation.

7. The process according to claim 1, wherein a hydrogenation mixture is withdrawn from the hydrogenation and subjected to a thermal work-up to obtain an alcohol-rich fraction, a low-boiler fraction and a high-boiler fraction.

8. The process according to claim 1, wherein the thermal separation unit comprises a thin film evaporator and a falling film evaporator, the thin film evaporator and the falling film evaporator being serially interconnected, optionally with the thin film evaporator being serially connected downstream of the falling film evaporator.

9. The process according to claim 1, wherein the thermal separation unit comprises two or three serially interconnected falling film evaporators.

10. The process according to claim 1, wherein the first membrane separation unit takes a form of a two-stage depleting cascade.

11. The process according to claim 1, wherein the catalyst system comprises a rhodium catalyst comprising an organophosphorus ligand, the ligand being selected from the group consisting of a phosphite, a phosphine and a phosphoramidite.

12. The process according to claim 1, wherein the olefin comprises eight carbon atoms and is hydroformylated to the aldehyde comprising nine carbon atoms, and the aldehyde is hydrogenated to an alcohol comprising nine carbon atoms.

13. A plant, comprising: a) at least one hydroformylation reactor comprising a reactant inlet and a product outlet; b) a first membrane separation unit comprising a first membrane entry point, a first permeate connection point and a first retentate connection point; c) a thermal separation unit comprising a product inlet, a head product connection point and a bottom product connection point; d) a second membrane separation unit comprising a second membrane entry point, a second permeate connection point and a second retentate connection point; e) at least one hydrogenation reactor comprising an aldehyde entry point and an alcohol exit point; wherein the product outlet of the hydroformylation reactor connects directly or via a devolatilizer to the first membrane entry point of the first membrane separation unit; the first retentate connection point of the first membrane separation unit connects to the reactant inlet of the hydroformylation reactor; the first permeate connection point of the first membrane separation unit connects directly or via a devolatilizer to the product inlet of the thermal separation unit; the bottom product connection point of the thermal separation unit connects to the second membrane entry point of the second membrane separation unit; the head product connection point of the thermal separation unit connects directly or via an adsorber to the aldehyde entry point of the hydrogenation reactor; the second retentate connection point of the second membrane separation unit connects together with the product outlet of the hydroformylation reactor to the first entry point of the first membrane separation unit; the second permeate connection point of the second membrane separation unit connects directly or via the adsorber to the aldehyde entry point of the hydrogenation reactor; and the second membrane separation unit takes a form of a two-stage enriching cascade.

14. The process according to claim 1, which is carried out in a plant comprising: a) the at least one hydroformylation reactor comprising a reactant inlet and a product outlet; b) the first membrane separation unit comprising a first membrane entry point, a first permeate connection point and a first retentate connection point; c) the thermal separation unit comprising a product inlet, a head product connection point and a bottom product connection point; d) the second membrane separation unit comprising a second membrane entry point, a second permeate connection point and a second retentate connection point; e) at least one hydrogenation reactor comprising an aldehyde entry point and an alcohol exit point; wherein the product outlet of the hydroformylation reactor connects directly or via a devolatilizer to the first membrane entry point of the first membrane separation unit; the first retentate connection point of the first membrane separation unit connects to the reactant inlet of the hydroformylation reactor; the first permeate connection point of the first membrane separation unit connects directly or via a devolatilizer to the product inlet of the thermal separation unit; the bottom product connection point of the thermal separation unit connects to the second membrane entry point of the second membrane separation unit; the head product connection point of the thermal separation unit connects directly or via an adsorber to the aldehyde entry point of the hydrogenation reactor; the second retentate connection point of the second membrane separation unit connects together with the product outlet of the hydroformylation reactor to the first entry point of the first membrane separation unit; and the second permeate connection point of the second membrane separation unit connects directly or via the adsorber to the aldehyde entry point of the hydrogenation reactor.

Description

(1) Further preferred embodiments of the invention will become apparent from the following detailed description of a plant according to the present invention and of the process according to the present invention which is carried out using a plant according to the present invention:

(2) FIG. 1 shows a flow diagram of a first embodiment comprising two separate adsorbers,

(3) FIG. 2 shows a flow diagram of a second embodiment of an inventive plant comprising a conjoint adsorber,

(4) FIG. 3 shows a detailed depiction of the purifying stage,

(5) FIG. 4 shows a first embodiment of the thermal separation unit, consisting of a falling film evaporator and a thin film evaporator,

(6) FIG. 5 shows a second embodiment of the thermal separation unit, consisting of two falling film evaporators,

(7) FIG. 6 shows a detailed depiction of a first membrane separation unit as a two-stage depleting cascade, and

(8) FIG. 7 shows a detailed depiction of a second membrane separation unit as a two-stage enriching cascade.

(9) FIG. 8 shows a simplified model of organophilic nanofiltration for calculating membrane separation.

(10) FIG. 9 shows a depiction of a membrane separation cascade for Variant A according to an embodiment of the invention.

(11) FIG. 10 shows a depiction of a membrane separation cascade for Variant B according to an embodiment of the invention.

(12) FIG. 11 shows a depiction of a membrane separation cascade for Variant C according to an embodiment of the invention.

(13) FIG. 12 shows a depiction of a membrane separation cascade for Variant D according to an embodiment of the invention.

(14) FIG. 13 shows a comparison of results of simulation computations with respect to the Rh consumption factor.

(15) FIG. 14 shows a comparison of results of simulation computations with respect to the high boiler concentration in the retentate.

(16) FIG. 15 shows a comparison of results of simulation computations with respect to the membrane area requirements.

(17) FIG. 16 shows a comparison of results of simulation computations with respect to the rhodium concentration in the separation unit.

(18) FIG. 1 shows the flow diagram of an inventive plant with which the inventive process can be carried out. The flow diagram is simplified for clarity, as usual. Self-evident plant components such as valves, pumps and the like are not depicted.

(19) At the heart of the plant is a hydroformylation reactor 1. This is where the hydroformylation reaction takes place. An olefin 2 is reacted with syngas 3a mixture of carbon monoxide and hydrogenin the presence of a homogeneously dissolved catalyst to corresponding aldehydes having one carbon atom more. This reaction is a gas/liquid phase reaction wherein the olefin and the reaction products are in the liquid phase, while one portion of syngas 3 forms the gaseous phase and another portion of the syngas is dissolved in the liquid phase. A homogeneous catalyst complex is likewise dissolved in the liquid phase.

(20) Optionally, a solvent can be supplied to the hydroformylation reactor 1, for example alkanes to accompany the olefin used. The hydroformylation then takes place in the presence of the optional solvent.

(21) Any type of reactor design which permits a gas-liquid phase reaction is possible in principle. A bubble column reactor is used with preference. Bubble column reactors are general common knowledge in the prior art and are described at length in ULLMANN: Deen, N. G., Mudde, R. F., Kuipers, J. A. M., Zehner, P. and Kraume, M.: Bubble Columns. Ullmann's Encyclopedia of Industrial Chemistry. Published Online: 15 Jan. 2010. DOI: 10.1002/14356007.b04_275.pub2

(22) Since bubble column reactors are not infinitely scalable owing to their flow behaviour, two or more comparatively small reactors connected in parallel have to be provided instead of a single large reactor for a plant designed to have a very large manufacturing capacity. A world-scale plant with a rating of 30 t/h may have two or three bubble columns each of 15 t/h or, respectively, 10 t/h capacity. The reactors operate in parallel under the same reaction conditions. Connecting two or more reactors in parallel also has the advantage that the reactor does not have to be run in the energetically unfavourable partial-load range when plant utilization is lower. Instead, one of the reactors is switched off completely and the other reactor continues to be run under full load. A triple arrangement can accordingly respond even more flexibly to demand changes.

(23) Any reference herein to a hydroformylation reactor is thus not necessarily to be understood as meaning one apparatus. Two or more mutually interconnected reactors may also be intended.

(24) The reaction is carried out under customary conditions.

(25) A temperature of 120? C. to 160? C. and a pressure of 20 to 28 MPa are particularly preferred. A conversion of >90% is sought under these conditions. Syngas having a hydrogen/carbon monoxide ratio of 1:1 is fed into the reactor in excess.

(26) Any olefin amenable to the oxo reaction is useful in principle as substrate for the hydroformylation. These are the olefins having two to twenty carbon atoms in particular. C6-C12 Olefin mixtures are used with particular preference. INA production utilizes olefin mixtures having a high isooctene content, for example the di-n-butene (CAS number 10071-47-9) available from Evonik Industries.

(27) The following homogeneous catalysts are useful as catalyst system:

(28) Rhodium-phosphite systems are used as homogeneous catalyst in particular. Rhodium nonanoate and tris(2,4-di-tert-butylphenyl)phosphite is an example of such a system. Metal concentration should be between 5 and 100 ppm and the ligand/rhodium ratio should preferably be about 5:1.

(29) The hydroformylation produces not only the desired aldehydes but also in a consecutive secondary reaction, the corresponding alcohols and also high boilers. The high boilers include inter alia dimers, trimers, aldol products, Tishchenko products, esters and ethers. High-boiler formation in the reaction is unwanted, since it leads to yield losses, but is technically impossible to fully avoid. High boilers therefore have to be removed from the system at a rate commensurate with their rate of formation. High boilers are so called because these substances have a higher boiling point than the aldehyde, and so high boilers collect at the base of the thermal separation unit and/or of the distillation downstream of the hydrogenation. By contrast, low boilers include olefins, alkanes and aldehydes formed in the hydroformylation or hydrogenation or already present in the olefin mixture.

(30) Reactor 1 has withdrawn from it a liquid hydroformylation effluent 4 which as well as the desired aldehyde also contains unconverted olefin, syngas dissolved in the liquid, the homogeneously dissolved catalyst system, further low boilers and the high boilers. Any optional solvent used forms part of the low boilers.

(31) The hydroformylation effluent 4 is cooled down to a temperature of about 40 to 50? C. in a first heat exchanger 5. The hydroformylation effluent 4 is decompressed to about 0.5 MPa in a first devolatilizer 9a, the effervescing syngas 3 being returned into reactor 1. The hydroformylation effluent 4 is then applied to a first membrane separation unit 6. The membrane separation unit 6 comprises a multistage depleting cascade, which is more particularly elucidated with reference to FIG. 6. For the purposes of understanding the functional interrelationships, however, it is sufficient to view the first membrane separation unit 6 as a single membrane.

(32) The incoming hydroformylation effluent 4 in the first membrane separation unit 6 is separated therein into a product stream 7 and a reactor return stream 8, while the catalyst system in hydroformylation effluent 4 partitions into the reactor return stream 8. The product stream 7 is the permeate of the first membrane separation unit 6, while the reactor return stream 8 forms the retentate of the first membrane separation unit 6. Since the first membrane separation unit 6 allows the catalyst system to pass at a distinctly lower rate than the other constituents of the hydroformylation effluent 4, the catalyst system collects in the reactor return stream 8. The membrane separation unit 6 is preferably operated such that about three-quarters of the catalyst system removed from the hydroformylation reactor 1 ends up in the reactor return stream 8. The first membrane separation unit thus has a 75% retention with regard to the catalyst system. The following operating parameters must be observed for this:

(33) Membrane temperature is between 20 and 160? C., preferably between 25 and 90? C. and more preferably between 30 and 60? C. A 10 to 30 K higher temperature can be advantageous in the concentrate loops. Transmembrane pressure difference is between 1 and 10 MPa, preferably between 1.5 and 5 MPa. It is particularly preferable to operate the membrane at about 2.5 to 3.5 MPa transmembrane pressure. The spiral-wound element is the membrane module design which is used with preference.

(34) Preference is given to using membranes having a separation-active layer composed of a material selected from cellulose acetate, cellulose triacetate, cellulose nitrate, regenerated cellulose, polyimides, polyamides, polyetheretherketones, sulphonated polyetheretherketones, aromatic polyamides, polyamideimides, polybenzimidazoles, polybenzimidazolones, polyacrylonitrile, polyaryl ether sulphones, polyesters, polycarbonates, polytetrafluoroethylene, polyvinylidene fluoride, polypropylene, terminally or laterally organomodified siloxane, polydimethylsiloxane, silicones, polyphosphazenes, polyphenyl sulphides, polybenzimidazoles, 6.6 Nylon?, polysulphones, polyanilines, polypropylenes, polyurethanes, acrylonitrile/glycidyl methacrylate (PANGMA), polytrimethylsilylpropynes, polymethylpentynes, polyvinyltrimethylsilane, polyphenylene oxide, alpha-aluminas, gamma-aluminas, titanias, silicas, zirconias, silane-hydrophobicized ceramic membranes as described in EP 1 603 663 B1, polymers with intrinsic microporosity (PIM) such as PIM-1 and others, as described for example in EP 0 781 166 and in Membranes by I. Cabasso, Encyclopedia of Polymer Science and Technology, John Wiley and Sons, New York, 1987. The abovementioned chemistries can be in crosslinked form, optionally as a result of the addition of co-chemistries, in the separation-active layer in particular, or, as mixed matrix membranes, be provided with fillers such as, for example, carbon nanotubes, metal organic frameworks or hollow spheres and also particles of inorganic oxides or inorganic fibres, for example ceramic or glass fibres.

(35) Particular preference is given to using membranes where the separation-active layer is a polymer layer composed of terminally or laterally organomodified siloxane, polydimethylsiloxane or polyimide which are constructed from polymers with intrinsic microporosity (PIM) such as PIM-1, or wherein the separation-active layer is built over a hydrophobicized ceramic membrane.

(36) Very particular preference is given to using membranes composed of terminally or laterally organomodified siloxanes or polydimethylsiloxanes. Membranes of this type are commercially available.

(37) In addition to the abovementioned materials, the membranes may comprise further materials. More particularly, the membranes may comprise scaffolding or support materials whereto the separation-active layer is applied. In composite membranes of this type, a scaffolding material is present alongside the actual membranes. A selection of scaffolding materials are described in EP 0 781 166, hereby incorporated herein by reference.

(38) A selection of commercially available solvent-stable membranes are the MPF and Selro series from Koch Membrane Systems, Inc., different types of Solsep BV, the Starmem? series from Grace/UOP, the DuraMem? and PuraMem? series from Evonik Industries AG, the Nano-Pro series from AMS Technologies, the HITK-T1 from IKTS, and also oNF-1, oNF-2 and NC-1 from GMT Membrantechnik GmbH and the Inopor?nano types from Inopor GmbH.

(39) The retentate of the first membrane separation unit 6referred to herein as reactor return stream 8 or else as primary recyclatecontains not only the high rhodium concentration but also the other chemistries of the hydroformylation effluent, namely aldehyde, olefin, dissolved syngas, further low boilers and high boilers. The reactor return stream 8 is returned into the hydroformylation reactor 1. The reactor return streamcontrary to the simplifying drawingneed not be fed into reactor 1 together with fresh olefin 2 and fresh syngas 3. It is perfectly conceivable for these three streams to be fed separately into the hydroformylation reactor 1 at different places.

(40) In order that the catalyst system may not lose its activity in the first membrane separation unit 6, this membrane separation step is conducted by maintaining a minimum CO partial vapour pressure. It should be, as described in EP1931472B1, at least 100 kPa. Therefore, decompression in devolatilizer 9a is not complete, but only down to 0.5 MPa. This is because the dissolved syngas is only supposed to be removed downstream of the membrane. For this, product stream 7 is decompressed in a second devolatilizer 9b to atmospheric pressure. The syngas remaining in the permeate of the first membrane separation unit escapes completely in the process and is removed from the plant.

(41) The devolatilized product stream 7 is then transferred into a thermal separation unit 10. This is most simply a distillation column, but preferably is a combination of a thin film evaporator and a falling film evaporator (cf. FIG. 4) or a combination of two or three falling film evaporators (cf. FIG. 5).

(42) Product stream 7 is evaporated in the thermal separation unit 10 by the action of heat. For this, the temperature at the base of the falling film evaporator is set to 90? C.; the temperature at the base of the thin film evaporator is 100? C. Evaporation is supported by an applied vacuum of about 30 hPa in both cases. In this way, more than 90% of the mass introduced into the thermal separation unit 10 with product stream 7 is evaporated. This vaporous mass forms the head product 11 of the thermal separation unit. Since the introduced components have different boiling points it is not just a purely quantitative separation of product stream 7 which is brought about by the evaporation but also a qualitative one: Aldehyde, alcohol and the other low boilers preferentially partition into head product 11 because of their lower boiling points. The unevaporated components form a liquid bottom product 12 consisting essentially of high boilers, aldehyde and catalyst system, wherein aldehyde and high boilers account for approximately the same weight fraction. The operating conditions of the thermal separation unit are so chosen that preferably 95% of the aldehydes introduced with product stream 7 end up in head product 11. Not more than 5% of the aldehydes introduced with product stream 7 remain in bottom product 12.

(43) The large stream withdrawn as top product 11, which contains the products of value, is then run through a first adsorber 13. Adsorber 13 has the function to catch residual quantities of catalyst, especially noble metal, entrained in droplets with the vapour. This is accomplished by using a conventional adsorbent such as activated carbon, silicates or aluminas, which are used in the form of a fixed bed. The adsorption is conducted at a temperature between 30 and 140? C. and space velocities of 0.01 to 5 1/h.

(44) Head product 11, now completely purified of its catalyst load, is then subjected to a hydrogenation 14. The hydrogenation takes place in two serially connected hydrogenation reactors. The first reactor is operated in loop mode, the second in straight-path mode. Hydrogenation takes place in the liquid phase in a temperature range from 120 to 220? C., preferably at a temperature of 160? C. under adiabatic conditions. Pressure is from 1.5 to 30 MPa. Hydrogenation is effected in a heterogeneous fixed-bed catalyst such as, for example, copper, cobalt, copper-nickel, copper-chromium, copper-chromium-nickel, zinc-chromium or nickel-molybdenum catalysts, which may optionally include still further elements. Details for designing a suitable hydrogenation are described in EP0987240B1 and EP0987241B1, and also in DE102008007080A1.

(45) Hydrogenation 14 exits into hydrogenation mixture 15 comprising essentially alcohol, alkanes (from unconverted olefins) and also hydrogenated low boilers and high boilers. Hydrogenation mixture 15 is then sent to a thermal work-up 16 and split therein into an alcohol-rich fraction 17, a low-boiler fraction 18 and a high-boiler fraction 19. The alcohol-rich fraction 17 is the actual product of value of the process according to the present invention. The low boilers 18 and high boilers 19 are by-products which can be marketed for subordinate purposes. The thermal work-up 16 of the three fractions 17, 18 and 19 from hydrogenation mixture 15 will be further elucidated with reference to FIG. 3.

(46) In the event that the conversion of alkenes in the hydroformylation is incomplete, head product 11 of thermal separation unit 10 will contain a sizeable quantity of unconverted alkenes. In order that these may not be lost in hydrogenation 14, the hydrogenation may also be placed downstream of the thermal work-up, specifically one hydrogenation for the alcohol-rich fraction 17 (which in this case is more aldehyde-rich) and one hydrogenation for the high boilers fraction 19. The low boilers fraction 18 then contains the unconverted alkenes and can be returned into the hydroformylation reactor 1 (as is not shown in this figure). In the event of an alkene recycle, the head product 11 is thus fed into the hydrogenation in part only, not as a whole; the alkenes are first separated off and recycled.

(47) As mentioned, bottom product 12 of thermal separation unit 10 contains essentially the high boilers, minor amounts of aldehyde and catalyst. The mass flow of bottom product 12 is distinctly less than that of the head product. When the product stream is 30 tonnes per hour and the proviso that 90% of introduced mass departs the thermal separation unit 10 overhead is observed, then the mass flow of bottom product 12 is merely 3 t per hour, i.e. 1/9 of that of the head product.

(48) Bottom product 12 is then applied to a second membrane separation unit 20. Bottom product 12 is therein separated into a permeate 21 and a retentate 22, with the catalyst system partitioning into the retentate 22, since the second membrane separation unit 20 retains the catalyst system preferentially. Owing to the low mass stream with which the second membrane separation unit 20 has to cope compared with the first membrane separation unit 6, the catalyst in bottom product 12 can be nearly completely retained and collected in retentate 22. This is accomplished particularly when the choice of membrane material is particularly permeable for high boilers and hence passes the high boilers into permeate 21. Retentate 22 then consists essentially of aldehyde and catalyst.

(49) The separation in the second membrane separation unit 20 is effected at a temperature between 20 and 160? C., preferably between 25 and 90? C. and more preferably between 30 and 60? C. The transmembrane pressure difference is between 1 to 10 MPa, preferably between 1.5 and 5 MPa. It is particularly preferable to operate the membrane at about 2.5 to 3.5 MPa transmembrane pressure. The spiral-wound element is the membrane module design which is used with preference.

(50) The first and second membrane units can use the same or different membrane materials.

(51) The classes of materials described hereinbelow are useful as membrane material for the second membrane separation unit 20:

(52) Preference is given to using membranes within the second membrane separation unit having a separation-active layer composed of a material selected from cellulose acetate, cellulose triacetate, cellulose nitrate, regenerated cellulose, polyimides, polyamides, polyetheretherketones, sulphonated polyetheretherketones, aromatic polyamides, polyamideimides, polybenzimidazoles, polybenzimidazolones, polyacrylonitrile, polyaryl ether sulphones, polyesters, polycarbonates, polytetrafluoroethylene, polyvinylidene fluoride, polypropylene, terminally or laterally organomodified siloxane, polydimethylsiloxane, silicones, polyphosphazenes, polyphenyl sulphides, polybenzimidazoles, 6.6 Nylon, polysulphones, polyanilines, polypropylenes, polyurethanes, acrylonitrile/glycidyl methacrylate (PANGMA), polytrimethylsilylpropynes, polymethylpentynes, polyvinyltrimethylsilane, polyphenylene oxide, ?-aluminas, ?-aluminas, titanias, silicas, zirconias, silane-hydrophobicized ceramic membranes as described in EP 1 603 663 B1, polymers with intrinsic microporosity (PIM) such as PIM-1 and others, as described for example in EP 0 781 166 and in Membranes by I. Cabasso, Encyclopedia of Polymer Science and Technology, John Wiley and Sons, New York, 1987. The abovementioned chemistries can be in crosslinked form, optionally as a result of the addition of co-chemistries, in the separation-active layer in particular, or, as mixed matrix membranes, be provided with fillers such as, for example, carbon nanotubes, metal organic frameworks or hollow spheres and also particles of inorganic oxides or inorganic fibres, for example ceramic or glass fibres.

(53) Particular preference is given to using membranes where the separation-active layer is a polymer layer composed of terminally or laterally organomodified siloxane, polydimethylsiloxane or polyimide which are constructed from polymers with intrinsic microporosity (PIM) such as PIM-1, or wherein the separation-active layer is built over a hydrophobicized ceramic membrane.

(54) A detailed description of such membranes for use in high-boiler removal is found in EP2401078A1.

(55) Very particular preference is given to using membranes composed of terminally or laterally organomodified siloxanes or polydimethylsiloxanes. Membranes of this type are commercially available.

(56) In addition to the abovementioned materials, the membranes may comprise further materials. More particularly, the membranes may comprise scaffolding or support materials whereto the separation-active layer is applied. In composite membranes of this type, a scaffolding material is present alongside the actual membranes. A selection of scaffolding materials are described in EP0781166, hereby incorporated herein by reference.

(57) A selection of commercially available solvent-stable membranes are the MPF and Selro series from Koch Membrane Systems, Inc., different types of Solsep BV, the Starmem? series from Grace/UOP, the DuraMem? and PuraMem? series from Evonik Industries AG, the Nano-Pro series from AMS Technologies, the HITK-T1 from IKTS, and also oNF-1, oNF-2 and NC-1 from GMT Membrantechnik GmbH and the Inopor?nano types from Inopor GmbH.

(58) The second membrane separation unit 20 is configured as a multistage enriching cascade. This membrane arrangement will be more particularly elucidated with reference to FIG. 7. To understand the function of second membrane separation unit 20 it is sufficient to assume that a single membrane is concerned.

(59) The second membrane separation unit 20 has withdrawn from it the retentate 22 which is cooled down to about 40 to 50? C. in a heat exchanger 23 and then mixed with the likewise cooled-down hydroformylation effluent 4 and returned into the first membrane separation unit 6. Returning the secondary recyclate (retentate 22) into the first membrane separation unit 6 at a point upstream of hydroformylation reactor 1 offers the decisive advantage of reducing interference between the control of the second membrane separation unit 20 and that of hydroformylation reactor 1. It also stops aldehyde being unnecessarily passed back into the reaction together with retentate 22 and reducing the yield of said reaction. The catalyst constituents returned via the secondary return stream 22 are very largely rejected again by the first membrane separation unit 6 and fed back into reactor 1 via the primary reactor return stream 8.

(60) The second membrane separation unit 20 provides the permeate 21, which very largely consists of high boilers and residual aldehyde and is passed through a second adsorber 24 to trap and secure residual quantities of catalyst. To remove noble catalyst metals from the liquid permeate 21, the adsorption is carried out at a temperature of 30 to 140? C. and space velocities of 0.01 to 5 1/h. The adsorbent is preferably used in the form of a fixed bed.

(61) Useful adsorbents include particularly activated carbon, surface-rich polysilicic acids such as silica gels (silicic xerogels), finely divided silica, surface-rich aluminas and alumina hydrates as well as spent or virgin (hydrogenation) catalysts.

(62) Chemically modified silica materials as disclosed in WO 2006013060 A1 have been found to be particularly advantageous adsorbents. Adsorbents of this type are available under the article name of Mercaptoalkyl-modified Silica, Type Rh H3, Batch No. 09-S26-001 from PhosphonicS Ltd, 114 Milton Park, Abingdon, OXON, OX14 4SA, United Kingdom.

(63) Permeate 21, now completely purified of catalyst residues by adsorption, is then fed together with the likewise adsorptively purified head product 11 to hydrogenation 14. Alternatively, it would be conceivable to feed head product 11 and permeate 21 into separate hydrogenation reactions (not depicted) instead of a conjoint hydrogenation 14.

(64) FIG. 2 shows a version of the plant of FIG. 1, where permeate 21 of the second membrane separation unit 20 and head product 11 of thermal separation unit 10 are run through a conjoint adsorber 25 and then subjected to hydrogenation 14. The use of a conjoint adsorber 25 makes it possible for less adsorbent to be used, lowering the operating costs of the plant. The operating conditions and the adsorbent in this version are chosen as just described regarding the second adsorber 24.

(65) FIG. 3 shows the thermal work-up 16 in detail. It consists of a serial arrangement of three distillation columns 26, 27 and 28, which are operated at atmospheric pressure or at reduced pressure. The hydrogenation mixture 15 is fed into the first column 26. The hydrogenation mixture is separated therein into a low-boiler fraction 18, which is withdrawn overhead, and a bottom fraction 29 consisting essentially of high boilers and alcohol. The first distillation column 26 has from 20 to 70, preferably from 28 to 65 theoretical plates. The temperature in the first distillation column 26 is preferably adjusted such that the head temperature is in the range from 85 to 110? C., preferably in the range from 95 to 100? C. and the pot temperature is in the range from 175 to 200? C., preferably in the range from 185 to 193? C.

(66) The bottom fraction 29 of the first column 26 is fed into the second distillation column 27. The alcohol-rich fraction 17 is removed therein overhead. It is preferably more than 98% target alcohol. The bottom product 30 of the second distillation column 27 is a mixture of high boilers and residual alcohol. To perform this separating duty, the second distillation column 27 has from 8 to 35, preferably from 10 to 30 theoretical plates. The temperature in the second distillation column 27 is preferably adjusted such that the head temperature is in the range from 150 to 180? C., preferably in the range from 160 to 170? C. and the pot temperature is in the range from 180 to 205? C., preferably in the range from 185 to 195? C.

(67) The bottom product 30 of the second distillation column 27 is finally run into a third column 28 whose bottom product is the high-boiler fraction 19. Its head product comprises residual quantities of alcohol, which are mixed with the alcohol-rich fraction 17. The third distillation column 28 has from 15 to 35, preferably from 20 to 30 theoretical plates. The temperature in the third distillation column is preferably adjusted such that the head temperature is in the range from 95 to 120? C., preferably in the range from 100 to 110? C. and the pot temperature is in the range from 160 to 190? C., preferably in the range from 165 to 175? C.

(68) The three fractions 17, 18 and 19 are removed from the system and marketed.

(69) FIG. 4 shows the internal construction of thermal separation unit 10. It is formed by two serially connected thermal separation apparatuses, namely a falling film evaporator 31 and a thin film evaporator 32. The falling film evaporator 31 is of conventional technical design. The liquid product stream 7 enters at the top of the falling film evaporator 31 and is distributed from there over a multiplicity of vertically extending down pipes 33. The down pipes 33 are surrounded by a heating jacket 34 heated with medium pressure vapour. The medium pressure vapour is water vapour used as heating medium which does not react with the process chemicals. Its pressure is between 1.2 and 2.4 MPa, depending on site conditions. The medium pressure vapour enters heating jacket 34 through a steam inlet 35, passes its heat via the walls of down pipes 33 to product stream 7 and exits again via a steam outlet 36, having cooled down. The liquid product stream 7 passes downwardly through the down pipes 33 in line with the force of gravity and in the course of its passage is heated up by the hot steam (about 120? C.). At the point of exit at the base of the down pipes 33, the components of the product stream 7 which boil at 120? C. are very largely evaporated. It must be borne in mind here that a negative pressure of 3 and 500 hPa prevails in falling film evaporator 31. The evaporated fractions of product stream 7 depart the falling film evaporator 31 via a gas exit 37. The components which have not evaporated collect in the bottom product 38 and pass from there into the thin film evaporator 32.

(70) The thin film evaporator 32 is similar to the falling film evaporator 31 in having a medium pressure vapour heated heating jacket 34 wherethrough the process steam flows in through a steam inlet 35 and departs again through a steam outlet 36, having cooled down. The steam heats a beak 39 from the outside, the inside surface of which is a support for the hitherto unvaporized fractions of product stream 7 from bottom product 38 of falling film evaporator 31. A rotor 40 is arranged coaxially within the beak 39 and turns about the longitudinal axis of thin film evaporator 32. It is equipped with a multiplicity of wipers 41, which spread the liquid feed into a thin film on the inside surface of beak 39. The fractions which evaporate in the process depart the thin film evaporator 32 via a gas exit 42 and are then combined with the evaporated components from falling film evaporator 31 (ex gas exit 37) to form the head product 11 of the thermal separation unit 10. In this way, about 90% of the mass introduced into the thermal separation unit 10 with product stream 7 is evaporated and withdrawn as head product 11.

(71) The remaining 10% of the introduced product stream 7 depart the thermal separation unit 10 in liquid form, namely from the base 43 of the thin film evaporator, where the fractions of the feed of the thin film evaporator 32 which have not evaporated within beak 39 collect. Bottom product 43 accordingly corresponds to bottom product 12 of thermal separation unit 10.

(72) FIG. 5 depicts an alternative embodiment of thermal separation unit 10. It consists of two serially connected falling film evaporators 31 and 44. The two falling film evaporators 31 and 44 correspond to the falling film evaporator 31 shown in FIG. 4 and therefore need not be further elucidated. Their respective gas exits 37 are combined to form the head product 11 of thermal separation unit 10. The bottom product 12 of thermal separation unit 10 is withdrawn from the base 45 of the second falling film evaporator 44. The bottom product 38 of the first falling film evaporator 31 serves as feed to the second falling film evaporator 44. The same method can be used to serially connect three falling film evaporators (not depicted).

(73) FIG. 6 shows the in-principle construction of first membrane separation unit 6. The first membrane separation unit 6 is a two-stage depleting cascade. It is fed with a mixture of the devolatilized hydroformylation effluent 4 of the reactor and retentate 22 of the second membrane separation stage by means of a pump 46 of a first stage 47. The permeate of the first stage 47 corresponds to the resulting permeate of the first membrane separation unit 6 and hence to product stream 7 of the plant. The retentate 48 of the first stage 47 is applied to a second stage 49 without further increase in pressure. The retentate 8 of the second stage 49 corresponds to the resulting retentate of the first membrane separation unit 6 and is recycled as reactor return stream 8 to a point upstream of hydroformylation reactor 1. The permeate 50 of the second stage 49 is mixed with the feed of the first membrane separation unit and fed via pump 46 back to the first stage 47. The permeate 50 of the second stage 49 thus corresponds to the internal permeate recycle of the membrane separation unit configured as a depleting cascade.

(74) FIG. 7 shows the internal construction of the second membrane separation unit 20. It is configured as a two-stage enriching cascade. The feed for the second membrane separation unit 20 is the bottom product 12 of the thermal separation unit 10. It is compressed by a first pressure elevation pump 51 to a pressure of about 3 MPa and applied to a first stage 52. The retentate of the first stage 52 corresponds to the resulting retentate of the second membrane separation unit 20 and departs the second membrane separation stage 20 as retentate 22/secondary recyclate and is mixed with the devolatilized hydroformylation effluent 40 and returned in this form into the first membrane separation unit 6.

(75) The permeate 53 of the first stage 52 is again brought by a second pressure elevation pump 54 to a pressure of about 3 MPa in order that the transmembrane pressure of the first stage 52 may be equalized. The second stage 55 of the membrane separation system then ensues. The resulting permeate 21 corresponds to the resulting permeate 21 of the second membrane separation unit 20. It is subjected to adsorptive purification and then subjected to hydrogenation. The retentate 56 of the second stage 55 is mixed with the feed of the second membrane separation stage 20 (=bottom product 12) and returned via the first pressure elevation pump 51 into the first stage 52. The retentate 56 of the second stage thus constitutes the internal retentate recycle of the enriching cascade.

EXAMPLES

(76) Variants of working under the hydroformylation effluent in the manner of the present invention will now be compared by means of simulations. Simulation is the means of choice because of the complexity of the plant structure.

(77) The process under consideration is the production of C.sub.9 alcohols from C.sub.8 olefins.

(78) Model of Hydroformylation

(79) The simulation describes the hydroformylation of the C.sub.8 olefin mixture dibutene in simplified form via a formally kinetic approach. The following reactions were taken into account. The main reaction is the hydroformylation of dibutene with synthesis gas (CO+H2) to form the C.sub.9 aldehyde nonanal (INAL) as per reaction 1:
dibutene+CO+H2.fwdarw.INALReaction 1:

(80) The further reaction taken into account is descendant reaction 2, the hydrogenation of the aldehyde INAL to the alcohol isononanol (INA):
INAL+H2.fwdarw.INAReaction 2:

(81) A particular point of interest in simulating this hydroformylation process, which employs nanofiltration, is how possible high boilers build up in the catalyst circuit. But high-boiler formation involves a multiplicity of unknown reactions. To keep the reaction system as simple as possible, therefore, the kinetic model only takes account of one further reaction to model high-boiler formation. Accordingly, there is only one high-boiling component in the simulation to represent the actual high-boiler mixture formed in the course of the hydroformylation. The high boilers are represented by dinonyl ether (DiEther) in the simulation. Dinonyl ether is formed as per reaction 3 from nonanal (INAL) and nonanol (INA):
INA+INAL+H2.fwdarw.H2O+DiEtherReaction 3:

(82) In principle, the selection of the high-boiler reaction is arbitrary. The ether formation in reaction 3 can thus also be replaced by some other reaction in which no water (H2O) is formed, an acetal formation for example.

(83) To map the dependencies of the reactions on the various quantities, the following equations were employed to model the reaction rates (in kmol m.sup.?3 min.sup.?1) r.sub.i, i=1, . . . , 3:

(84) $\begin{array}{cc}{r}_1=c_{\mathrm{total}}{k}_1\left(x_{\mathrm{dibutene}}^{n_1}-{\left(\frac{x_{\mathrm{nonanal}}}{k_{\mathrm{ggw}}{p}^2}\right)}^{n_1}\right)c_{\mathrm{Rh}}^{n_{\mathrm{Rh}}}{k}_{\mathrm{Li}}{k}_p& \left(1\right)\\ r_2=c_{\mathrm{total}}{k}_2{x}_{\mathrm{nonanal}}{c}_{\mathrm{Rh}}^{n_{\mathrm{Rh}}}& \left(2\right)\\ r_3=c_{\mathrm{total}}{k}_3{x}_{\mathrm{nonanal}}{x}_{\mathrm{nonanol}}& \left(3\right)\end{array}$
where c.sub.total is the total amount of substance concentration [kmol/m.sup.3], x.sub.i is the molar fraction of component i, p is the pressure in bar, and c.sub.Rh, is the Rh concentration in ppm. The dependence on the Rh concentration is mapped via the exponent nRh; n.sub.1 is the order of the first reaction. The reaction rate k.sub.i is modelled using the Arrhenius approach:

(85) $\begin{array}{cc}{k}_i=k_{0,i}\exp \left(-\frac{E_{\mathrm{Ai}}}{\mathrm{RT}}\right)& \left(4\right)\end{array}$

(86) The term k.sub.Li is used to represent the dependence of the hydroformylation reaction R1 on the ratio between ligand and rhodium.

(87) $\begin{array}{cc}{k}_{\mathrm{Li}}=1+\frac{k_{\mathrm{Li},1}{X}_{\mathrm{Li}}}{1+k_{\mathrm{Li},2}{X}_{\mathrm{Li}}^2}& \left(5\right)\end{array}$

(88) X.sub.Li is the molar ratio between ligand and rhodium. The pressure dependence k.sub.p is represented by:
k.sub.p=tan h(k.sub.p,0p)(6)

(89) Finally, the constant k.sub.ggw is used to describe the (pseudo) equilibrium between dibutene and nonanal. The values of all the constants are summarised in Table 1.

(90) TABLE-US-00001 TABLE 1 Values of constant reaction parameters Parameter Value Parameter Value k.sub.0.1 13537 min.sup.?1 E.sub.A,1 56037 kJ/kmol K k.sub.0.2 26810 .Math. 10.sup.3 min.sup.?1 E.sub.A,2 92703 kJ/kmol K k.sub.0.3 1840 min.sup.?1 E.sub.A,3 50858 kJ/kmol K k.sub.Li,1 2.453 k.sub.Li,2 0.01342 k.sub.p,0 0.004975 bar.sup.?1 k.sub.ggw 0.0601 bar.sup.?2 n.sub.1 1.452 n.sub.Rh 0.6

(91) The hydroformylation reaction (R1) is actually not an equilibrium reaction. However, autoclave test results have shown that complete dibutene conversion is not attained by the end of a 6 h run. A possible explanation for this is that the less speedily reacting di-methylhexene isomers have still not been completely converted by the end of the runs. However, the simple model of the formal kinetics which is used here does not distinguish between the various dibutenes. Introduction of a pseudo equilibrium between dibutene and nonanal is a way to describe the incomplete conversion. The square pressure dependence in equation (1) follows from the pseudo equilibrium condition:

(92) $\begin{array}{cc}{k}_{\mathrm{ggw}}^{*}=\frac{x_{\mathrm{nonanal},\mathrm{ggw}}}{x_{\mathrm{dibutene},\mathrm{ggw}}{x}_{\mathrm{CO},\mathrm{ggw}}{x}_{H_2,\mathrm{ggw}}}?a\frac{x_{\mathrm{nonanal},\mathrm{ggw}}}{x_{\mathrm{dibutene},\mathrm{ggw}}{p}_{\mathrm{CO}}{p}_{H_2}}?a^{*}\frac{x_{\mathrm{nonanal},\mathrm{ggw}}}{x_{\mathrm{dibutene},\mathrm{ggw}}{p}^2},k_{\mathrm{ggw}}=\frac{k_{\mathrm{ggw}}^{*}}{a^{*}}& \left(7\right)\end{array}$

(93) Only the square pressure dependence of the equilibrium term allows for satisfactory kinetic modelling of the experimental results at varying pressure.

(94) Model of Organophilic Nanofiltration (Membrane Separation)

(95) Organophilic nanofiltration through a membrane is mapped by a simple model for the purposes of the simulation. In this model, the transmembrane flux is computed as a function of the temperature, of the transmembrane pressure and of the composition on the retentate side and on the permeate side. The simplified approach which forms the basis of the model does not compute any locally distributed concentration profile and neglects the pressure drop in the flow across the membrane. The assumption that the compositionand hence also the driving concentration differenceacross the full membrane area is the same as at the membrane exit point causes the model to underestimate the separation effect of the membrane and to overestimate the area. Owing to the simplicity of its equations, however, the membrane module is useful for an initial screening of the various versions of the process by simulation. A simplified model of membrane separation as organophilic nanofiltration is shown by FIG. 8. FIG. 8: Simplified model of organophilic nanofiltration

(96) The molar permeate flux n.sub.M,i of component i (see FIG. 8) is computed in the model via the pure component flux n.sub.M,i, pure:

(97) $\begin{array}{cc}{n}_{M,i,\mathrm{pure}}^{}=P_{i,p}^0.Math.\exp \left(-{?}_{i,p}?p\right)?p\frac{{?}_{i,\mathrm{pure}}\left(T_{\mathrm{Ref}}\right)}{{?}_{i,\mathrm{pure}}\left(T\right)M_i}& \left(8\right)\end{array}$

(98) Here Pi,p0 is the standard permeance (mass specific) of the membrane for component i at a transmembrane pressure of 0 bar. The parameter ?.sub.i,p describes the compacting of the membrane, ?p is the transmembrane pressure and ?.sub.i,pure is the viscosity of the pure material and Mi is its molar weight. The pure component flux and the molar volume {tilde over (V)}i is used to compute the permeance of the membrane,

(99) $\begin{array}{cc}{P}_{i,p}=n_{M,i,\mathrm{pure}}^{}/1-\left(\exp \left(\frac{-?p{\tilde{V}}_i}{\mathrm{RT}}\right)\right)& \left(9\right)\end{array}$

(100) The permeance can finally be used to determine the permeate flux n.sub.M,i:

(101) $\begin{array}{cc}{n}_{M,i}^{}=P_{i,p}.Math.\left(x_{R,i}-x_{P,i}\exp \left(\frac{-?p{\tilde{V}}_i}{\mathrm{RT}}\right)\right)& \left(10\right)\end{array}$

(102) Table 2 shows the standard permeances for the simulation:

(103) TABLE-US-00002 TABLE 2 Standard permeances Material Permeance ONF 2 Dibutene 0.074 kg h.sup.?1 m.sup.?2 Nonanal 0.140 kg h.sup.?1 m.sup.?2 Nonanol 0.150 kg h.sup.?1 m.sup.?2 High boiler 0.105 kg h.sup.?1 m.sup.?2 Complex 0.026 kg h.sup.?1 m.sup.?2 Ligand 0.099 kg h.sup.?1 m.sup.?2

(104) There are various factors affecting the economics of the process. For one, the rhodium consumption factor (corresponds to the loss of Rh) should be very low; for another, the capital costsinter alia dependent on the membrane area neededshould not be too high.

(105) In the variants under consideration, the overall volume of the reactor is 67 m.sup.3. Aspects such as heat transfer or the geometry of the reactor were left out of the modelling. The dibutene feed is 20 t/hr, so presuming an annual on-stream time of 8500 hours a yield of 93% for conversion into the product nonanal is needed to attain the world scale standard of 200 kt/a. This yield is unattainable with a dibutene residual content of 8%, so recycling of the unreacted dibutenes is required in this case. This is taken into account in the simulation computations regarding interconnection variants A to D. The membrane stages all employed the oNF2 from GMT Membrantechnik GmbH. Employing other membranes in the entire plant or else only in parts of the plant might further improve the economics.

(106) Membrane temperature is 33? C. in the simulation computations regarding interconnection variants A to D. Higher operating temperatures for the nanofiltration reduce the membrane area while at the same time the membrane retention for the catalyst system decreases faster over time. Depending on membrane replacement costs, higher operating temperatures may be more economical in order to reduce total installed membrane area. Transmembrane pressure difference for the simulations performed is 35 bar (3.5 MPa).

(107) Four operative variants A to D of the invention will now be more particularly investigated:

(108) Interconnection Variant A

(109) Owing to the high rhodium losses in the nonanal product stream in a purely membrane-based separation, the high boiler and the remaining rhodium are hereinbelow separated from the nonanal stream downstream of the first membrane separation stage (NF1) by a thermal separation unit in the form of a thin film evaporator (DSV). This is depicted in FIG. 9. FIG. 9: Interconnection variant A

(110) Owing to the small rhodium concentration in the permeate, the thermal separation leads to a small loss of rhodium due to clustering. This rhodium loss is left out of the simulation.

(111) As a result of the thermal separation, the high boilers build up in the catalyst circuit. A portion of the high boiler stream is therefore separated off via a second membrane separation unit (NF2) in order to avoid an excessive build-up in the concentration of high boiler in the catalyst circuit. The further processing of the high boiler export, which is of interest by reason of the still considerable nonanal concentration, is not further considered in the simulation. After removal of the high boiler, the unreacted dibutene, which still comprises ?7% of the product stream, is separated off and returned into the reactor in order to achieve full conversion for the dibutene and attain the required nonanal production of 200 kta. The thermal reprocessing of the product mixture is modelled in the simulation as a simple splitter featuring fixed splitting factors. The first separating step is a flash evaporator operated at 40 mbar. The evaporator setting is established to ensure that 98% of the high boiler remains in the bottom product.

(112) Table 3 shows the results of the simulation. The computed membrane areas are 2416 m.sup.2 and 384 m.sup.2 for the first membrane separation unit and the second membrane separation unit, respectively. The rhodium consumption factor is 0.145 g of rhodium per metric ton (t) of nonanal. Of this, 38.9% is removed via the permeate from the second membrane separation unit. It is simple to further reduce this fraction by means of an adsorber or a further nanofiltration stage. The remaining 61.1% are losses due to clustering and segregation within the plant. The bottom product of the thermal separation unit DSV in this interconnection has a lower rhodium concentration than the retentate of the first membrane separation unit. The proportion accounted for by clustering is comparatively high.

(113) TABLE-US-00003 TABLE 3 Results of interconnection variant A Permeate 1 Retentate 1 Bottoms Permeate 2 Rate [t/h] 30.1 1.2 2.8 0.4 Dibutene [wt %] 10.5% 6.4% 1.1% 0.8% Nonanal [wt %] 82.7% 47.3% 45.5% 51.4% Nonanol [wt %] 2.1% 1.3% 3.9% 4.4% High boiler [wt %] 4.7% 45.0% 49.5% 43.4% Rhodium [ppm] 12.2 222.1 130.9 3.3

(114) Interconnection Variant B

(115) In variant B, the bottom product of the thermal separation unit (DSV) is routed to a point upstream of the first single-stage membrane separation unit (NF1). The high boilers are exported by feeding the retentate of the first membrane separation unit (NF1) to the second two-stage membrane separation unit (enrichment cascade NF2). This interconnection is depicted in FIG. 10. FIG. 10: Interconnection variant B

(116) Table 4 shows the results of the simulation. The computed membrane areas are 3032 m.sup.2 and 392 m.sup.2 for the first membrane separation unit and the second membrane separation unit, respectively. The rhodium consumption factor is 0.144 g of rhodium per metric ton of nonanal. Of this, 40.3% is removed via the permeate from the second nanofiltration. It is simple to further reduce this proportion by means of a scavenger or a further nanofiltration stage. The remaining 59.7% are losses due to clustering and segregation within the plant.

(117) TABLE-US-00004 TABLE 4 Results of interconnection variant B Permeate 1 Retentate 1 Bottoms Permeate 2 Rate [t/h] 30.9 4 3.7 0.4 Dibutene [wt %] 9.7% 6.4% 1.0% 6.3% Nonanal [wt %] 80.8% 50.5% 36.2% 56.9% Nonanol [wt %] 2.2% 1.4% 3.3% 1.6% High boiler [wt %] 7.3% 41.6% 59.5% 35.3% Rhodium [ppm] 15.5 157 129.4 3.5

(118) Interconnection Variant C

(119) Interconnection variant C is depicted in FIG. 11. The retentate of the first membrane separation unit (NF1) and the bottom product of the thermal separation unit (DSV) are mixed at a point upstream of the second membrane separation unit (NF2) and this mixture is run into the second membrane separation unit for high boiler exportation. In order that a catalyst cycle of 4 t/h may continue to be maintained, the retentate rate of the first membrane separation unit was reduced to 1.2 t/h. The retentate of the second membrane separation stage is recycled to upstream of the reactor. FIG. 11: Interconnection variant C

(120) Table 5 shows the results of the simulation. The computed membrane areas are 2473 m.sup.2 and 388 m.sup.2 for the first nanofiltration and the second nanofiltration, respectively. The rhodium consumption factor is 0.152 g of rhodium per metric ton of nonanal. Of this, 42.1% is removed via the permeate from the second nanofiltration. It is simple to further reduce this proportion by means of a scavenger or a further nanofiltration stage. The remaining 57.9% are losses due to clustering and segregation within the plant.

(121) TABLE-US-00005 TABLE 5 Results of interconnection variant C Permeate 1 Retentate 1 Bottoms Permeate 2 Rate [t/h] 30.2 1.2 2.9 0.4 Dibutene [wt %] 10.5% 6.4% 1.1% 2.4% Nonanal [wt %] 82.4% 46.6% 44.3% 51.1% Nonanol [wt %] 2.1% 1.2% 3.8% 3.5% High boiler [wt %] 5.0% 45.8% 50.8% 43.0% Rhodium [ppm] 12.4 217.8 129.6 3.8

(122) Interconnection Variant D

(123) Interconnection variant D as depicted in FIG. 12 shows a mode wherein, as in the case of variant A, the bottom product of the thermal separation means (DSV) is fed to the second membrane separation unit (NF2) for the purpose of high boiler exportation. However, the retentate of the second membrane separation unit is mixed with the hydroformylation effluent and fed to the first membrane separation unit and not, as in the case of variant C, returned to a point upstream of the reactor. FIG. 12: Interconnection variant D

(124) Table 6 shows the results of the simulation. The computed membrane areas are 2324 m.sup.2 and 382 m.sup.2 for the first membrane separation unit (NF1) and the second membrane separation unit (NF2), respectively. The rhodium consumption factor is 0.138 g of rhodium per metric ton (t) of nonanal. Of this, 37.7% is removed via the permeate from the second membrane separation unit. It is simple to further reduce this fraction by means of an adsorber or a further nanofiltration stage.

(125) The remaining 62.3% are losses due to clustering and segregation within the plant. The concentration in the bottom product of the thermal separation unit is lower than in the retentate of the first membrane separation unit, leading to a reduced level of clustering. In addition, a larger retentate stream can be run in the first membrane separation unit than in the case of interconnections A and C.

(126) TABLE-US-00006 TABLE 6 Results of interconnection variant D Permeate 1 Retentate 1 Bottoms Permeate 2 Rate [t/h] 29.9 4 2.7 0.4 Dibutene [wt %] 10.1% 7.3% 1.1% 0.8% Nonanal [wt %] 83.3% 57.7% 47.4% 53.3% Nonanol [wt %] 2.2% 1.6% 4.1% 4.7% High boiler [wt %] 4.3% 33.4% 47.4% 41.2% Rhodium [ppm] 11.3 159.3 125.4 3.1

(127) Conclusion

(128) FIGS. 13 to 16 give a graphic juxtaposition of the results of simulated interconnection variants A to D. FIG. 13: Comparison of results of simulation computations in respect of Rh consumption factor; FIG. 14: Comparison of results of simulation computations in respect of the high boiler concentration in the retentate; FIG. 15: Comparison of results of simulation computations in respect of the membrane area requirements; FIG. 16: Comparison of results of simulation computations in respect of the rhodium concentration in the separation unit.

(129) Comparing the graphic depiction in FIGS. 13 to 16 shows that variant D is the most favourable one with respect to all the parameters relevant to the economics of product removal. Variant D allows the smallest rhodium losses due to exportation and clustering and also has the lowest membrane area requirements.

(130) Of all invention embodiments A to D, therefore, the interconnection variant Dcharacterized by returning the retentate of the second membrane separation stage to a point upstream of the first membrane separation stageis the preferred one.

LIST OF REFERENCE SIGNS

(131) 1 hydroformylation reactor 2 olefin 3 syngas 4 hydroformylation effluent 5 first heat exchanger 6 first membrane separation unit 7 product stream 8 reactor return stream/primary recyclate 9 devolatilizer 10 thermal separation unit 11 head product 12 bottom product 13 first adsorber 14 hydrogenation 15 hydrogenation mixture 16 thermal work-up 17 alcohol-rich fraction 18 low-boiler fraction 19 high-boiler fraction 20 second membrane separation unit 21 permeate 22 retentate/secondary recyclate 23 second heat exchanger 24 second adsorber 25 conjoint adsorber 26 first distillation column 27 second distillation column 28 third distillation column 29 bottom product of first distillation column 30 bottom product of second distillation column 31 falling film evaporator 32 thin film evaporator 33 down pipes 34 heating jacket 35 steam inlet 36 steam outlet 37 gas exit from falling film evaporator 38 base of falling film evaporator 39 beak 40 rotor 41 wipers 42 gas exit from thin film evaporator 43 base of thin film evaporator 44 second falling film evaporator 45 base of second falling film evaporator 46 pump of depleting cascade 47 first stage of depleting cascade 48 retentate of first stage of depleting cascade 49 second stage of depleting cascade 50 permeate of second stage of depleting cascade/permeate recycle 51 first pressure elevation pump of enriching cascade 52 first stage of enriching cascade 53 permeate of first stage of enriching cascade 54 second pressure elevation pump of enriching cascade 55 second stage of enriching cascade 56 retentate of second stage of enriching cascade/retentate recycle