DK174479B1 - Process for the preparation of isopropyl alcohol and tertiary C4-5 alcohols - Google Patents

Description

in DK 174479 B1 o

The present invention relates to a process for preparing isopropyl alcohol and tertiary Ci al alcohols by reacting a hydrocarbon stream containing propylene or C4 ^4 isoolefins with water in the presence of a strong acidic solid catalyst at elevated pressure and temperature, the discharge of the residual gas and the aqueous alcohol product, and the extraction of the alcohol.

A German method for the continuous production of isopropyl alcohol and diisopropyl ether by catalytic hydration of propylene is already known from German Patent Specification No. 1,210,768. Here, as a catalyst, a strong acid cation exchanger consisting of one having approx. 5 to approx.

20% by weight of divinylbenzene crosslinked styrene polymerate containing approx. 1 sulfonic acid group per aromatic ring.

15 In this known method of producing the alcohol as the main product, a pressure of approx. 17 to approx. 105 atmospheres at a temperature of approx. 135 to approx. 157 ° C and with a ratio of 4 to 10 moles of water per day. moles of propylene. Furthermore, filling rates of 0.5 to 20 10 parts by volume of liquid propylene per unit are expected. volume part wetted resin catalyst .and hour. Since the density of liquid propylene at 20 saturated pressure d ^ is 0.51934 g / ml, this filling rate corresponds to approx. 6.7 to 123.4 moles of propylene per catalyst to approx. 6.7 to 123.4 moles of propylene per per liter of ka-90 mole percent of the propylene used could be converted per liter. gene run, where a conversion rate of approx. 35% is preferred.

Under these conditions, the best selectivity for isopropyl alcohol, hereinafter referred to as IPA, is obtained at 135 ° C, however, IPA constitutes only 69 mole percent of propylene converted for only 22%. The by-product selectivity for diisopropyl ether is approx. 28 and for propylene polymerisates approx.

3 mole percent.

As can be seen from this disclosure, the river's relatively low degree of conversion of the olefin at higher working temperature increases somewhat, although po

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the lyophilization as well, and the selectivity for IPA goes even further. In addition, temperatures above approx.

149 ° C to be a disadvantage for the useful life of the catalyst.

In the known process, it is necessary as well as difficult to keep the temperature fluctuations in the catalyst layer within a range of approx. 11 ° C, preferably 5.5 ° C, especially at higher degrees of conversion, although attempts to remedy this by local overheating of the catalyst material caused difficulties at a relatively high water / -10 olefin mole ratio of approx. 4 to approx. 10: 1.

A similar process is described in German Patent Specification No. 1,105,403, which uses as a catalyst a sulfonated blend polymer of approx.

88 to approx. 94% styrene and 12 to 6% p-divinylbenzene containing 12 to 16% by weight sulfur in the form of sulfonic acid groups and in which 25 to 75% of the protons of these acid groups have been replaced by metals of groups I or VIII in the periodic system, especially Cu. In this process, a reaction temperature of 120 to 220, in particular 155 to 220 ° C, is indicated at a filling rate of 0.5 to 1.5 volumes of liquid olefin per liter. catalyst and hour volume, as well as a water / olefin mole ratio of 0.3 to 1.5. As the examples in this disclosure show, this method also achieves only an acceptable selectivity for lower temperature IPA, approx. 120 ° C and low conversion rate, approx. 3.9 mole percent. Admittedly, at higher temperature (170 ° C) the conversion of propylene increases to approx. 35 mole percent, at the same time, however, IPA selectivity drops to 55%, and this contains approx. 45% diisopropyl ether.

The known method is only economically portable due to its low selectivity for IPA when allowing the high ether content, ie. can add an application. Like especially. further disclosed in the known methods is the disadvantage of a relatively short useful life of the 35 strong acidic ion exchanger used as a catalyst.

3

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DK 174479 B1 German Patent No. 2,233,967 describes an improved process for the preparation of IPA, which is carried out in a reactor developed as a column of columns, where propylene and water are fed in co-flow. A disadvantage of this method is a conversion rate of only 75% of propylene, which necessitates a recovery of propylene from the remaining gas stream, and the selectivity which is only approx. 95%.

German Patent No. 2,429,770 discloses a process by which isopropyl alcohol is prepared in a reactor in which propylene and reaction water are fed into the sump and passed through it through the reactor. Admittedly, this method achieves low propylene turnover of approx. 10% per review, a selectivity of 99%, which, however, at a load of the supercritical gas phase by approx. 20% 15 IPA drops to below 95%. Here, too, there is a disadvantage with the poorer overall turnover, which makes a concentration of the approx. 85% of the remaining gas needed.

Also described is a method of improving the selectivity which, upon returning the formed ether to the input stream, shifts the equilibrium position, cf. eg. Chemical Enginerring, September 4, 1972, pages 50, 51 and German Publication No. 2,759,237.

Alternative options for increasing the overall selectivity of a continuous process consist in the separate cleavage of the ether formed to either starting olefin and alcohol, cf. eg. U.S. Patent No. 4,352,945, or in two molecules of alcohol per liter. moles of ether in the presence of an excess of water, cf. German Publication Publication No. 3,336,644.

Both of the latter options have, among other things, the disadvantage that a further reactor is needed for their design. In the former option, the feed back to the input stream provides a strong decrease in reactor performance.

German Patent No. 3,419,392 discloses a process by which the ether is introduced into the reactor.

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B1 prepared from the reactants 5 to 30% before the end of the reaction run, based on the total length. However, in this process, the ether for recovery must first be recovered distillatively or extractively from the crude alcohol mixture.

Various methods are also known for the preparation of especially tertiary alcohols by hydration of corresponding olefinic unsaturated compounds in the presence of cation exchange resins of the sulfonic acid type. Only about 10-fold yield is obtained only by adding a solvent as the solvent-promoting agent which ensures that water and sec-olefin form a uniform phase or that olefin is present in a higher concentration in the process water.

Mentioned, e.g. a process by which isobutene or a hydrocarbon stream containing isobutene is reacted with an aqueous solution of an organic acid in the presence of an acidic cation exchange resin as catalyst, cf. German Publication Publication No. 2,430,470; further discloses a process by which a reaction system is added to a monovalent alcohol and a catalyst is used as above, cf. Japanese Application No. 137,906 / 75. According to U.S. Patent No. 4,096,194, glycol, glycol ether or glycol diether is added, and Japanese Publication No. 59802/76 (Chemical Abstracts, Vo. 86, 1977, 86: 19392 25) describes the reaction in the presence of polar solvents, in particular dioxane.

German publication specification 3,029,739 proposes polyvalent alcohols of the neopentyl type as a solvent, and in German publication no.

Nos. 2,721,206 and 3,031,702 suggest sulfones, e.g. sulfolane.

These known processes for the preparation of tertiary alcohols, especially tert-butyl alcohol, by hydrating secolefins and especially isobutene have the disadvantage of forming by-products such as addition products from isobutene 35 and the added organic acids or organic solvents, although These improve the reaction rate to a certain extent. Since these by-products and added solvents have boiling points near or below the tert-butyl alcohol, the separation and isolation of the tert-butanol from these 5 by-products and solvents is somewhat difficult and requires high operating costs. While the process of German Publication No. 2,721,206 for the preparation of sec-butanol by reaction of butene-1 and butene-2 at a temperature of 100 to 220 ° C shows good stability of the solvent used, it is. impossible to generate tert-butanol in greater yield by selectively hydrating the iso-butene in this way because, in addition to reacting with n-olefins present, the isobutene also reacts with water, and besides tert-butyl alcohol, sec-butyl alcohol and isobutene dimers also appear in larger extent. When using organic acids, such as acetic acid, precautions must be taken to limit corrosion problems.

According to German Patent Application No. 1,176,114, processes which are carried out without solvents are proposed, however, they lead to inferior reactions or to inferior selectivities. 1 German publication no.

3,031,702, it can be seen from Table II, page 14, that without solubilizing agents no economically viable isobutene reactions are obtained.

SUMMARY OF THE INVENTION It is thus an object of the invention to provide a process which, by direct hydration of olefins without ether return and separate decomposition thereof, allows the production of alcohols with good yield and greater selectivity, and in particular to produce tertiary alcohols with good yields. and selectivities from isoole veneers without addition. Solvent-promoting agents and the associated disadvantages.

This object is achieved by the method according to the invention, which is characterized by the characterizing part of claim 1.

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6 DK 174479 B1

Here, either by the sump phase process, the hydrocarbon stream and the process water can be fed to the sump of the individual reactors, or by the rubbing process both the process streams lead to the top of the individual 5 reactors. Preferably, the process of the invention is carried out in a reactor system comprising 2 to 10, especially 3 to 5 reactors. ! Proceeding in the same manner, the process of the invention may also be carried out in a single reactor containing 10 multiple reactor layers which. separate reaction zones where one process stream is fed from above. to the bottom, and the second process stream is conducted from the bottom to the top, and when flowing through the individual reaction zones in the co-stream, the process streams are guided by the sump phase process to the sump of the individual zone and by the rubbing method to the top of the individual zone. Advantageously, the hydrocarbon stream, hereinafter referred to as "residual gas", leaves the reactor system for removing alcohol with water and is then used as process water or as part of the process water.

The process according to the invention is carried out under already known process conditions:

In the reaction of propylene, the temperature is .12Ch-200 ° C, preferably 130 to 170 ° C. The reaction pressure is 60-200 bar, preferably 80 to 120 bar. The mole ratio of water to propylene is 1 to 50: 1, preferably 10 to 30: 1. LHSV is 0.2 to 3 hours \ preferably 0.5 to 1 hour

In the reaction of C ^ is isoolefins, the temperature is 30 to 150 ° C, preferably 60-110 ° C. The reaction pressure is 30 to 50 bar, preferably 10 to 20 bar. The total LHSV

is 0.2 to 5 hours \ preferably 0.2 to 2 hours The molar ratio of water to isobutene is 20-150: 1, preferably 40 to 90: 1.

In the preparation of isopropyl alcohol according to the invention, without concentration of the residual gas in one pass, propylene conversion of up to 99.5% is obtained. The selectivity of the process is at 99.0 to 99.5%.

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7 DK 174479 B1

Thus, the present invention provides, inter alia, a process for producing isopropyl alcohol, wherein the process is reacted with propylene or a propylene-containing hydrocarbon mixture in the presence of a highly acidic cation exchange resin, especially one of the sulfonic acid type. The process is characterized in that the propylene hydrocarbon stream is fed to the sump of the first reactor and the process water is fed to the sump of the last reactor. In the individual reactors, the process water and hydrocarbon stream are co-flowed. However, based on the overall system of the reactors, the hydrocarbon stream is cascaded from the first reactor to the last reactor and in the countercurrent process process / second cascade is fed from the last to the first reactor. On the top of the first reactor is removed an aqueous isopropyl alcohol which is known to be distillatively worked up in known manner. Depending on the reaction conditions, the recovered isopropyl alcohol has a purity of 99 to 99.9%. It contains only minor proportions (0.1 to 0.2%) of diisopropyl ether (DIPE), based on 100% alcohol. The greater amount of ether 20 (DIPE) - approx. 0.4 to 0.6%, based on the amount of alcohol formed - is discharged from the reactor with the remaining gas stream. The remaining gas contains only a small proportion of propylene so that a recovery operation is no longer necessary.

The process may also be carried out as a method of chewing gum in the process of the invention. Here it is headed. the propylene-containing hydrocarbon stream to the top of the first reactor. The process water is conducted to the top of the last reactor and, after phase separation in the sump of this reactor, the process water is removed and it is directed again in the opposite direction of the hydrocarbon stream to the top of the next reactor. Then, in the sump of the first reactor, the aqueous crude alcohol is extracted and it is worked up distillatively. On the sump of the last reactor, a residual gas containing only 35 liters of propylene is removed.

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8 DK 174479 B1

The secondary alcohol formed by the reaction of C ^ isoolefins as secondary by-product can be removed as a side-stream by the azeotroping of the tertiary alcohol from the pre-dewatering column. In particular, the side removal takes place at a few bottoms, preferably 1 to 7 bottoms, above the aqueous alcohol initiation bottom.

It has now been found that even without the addition of solvents, large isobutene conversions and large selectivities can be obtained, and in the method according to the invention, isobutene conversions in excess of 98% can be obtained. The selectivity of the process is at 99 to 99.9%. The raffinate leaving the last reactor is practically free of alcohol after washing with water.

As with the reaction with propylene, the process according to the invention has the advantage over the known processes also in the reaction of C ^ is isoolefins, because of the olefin content hydrocarbon mixtures, that the olefin is also reacted. almost full-20 constantly.

Thus, the present invention provides a process for the production of alcohols, especially isopropyl alcohols (IPA) and tert-butyl alcohol (TBA), by which process an olefin-containing hydrocarbon stream, preferably propylene or isobutylene, or a hydrocarbon stream containing these defines are reacted with water in the presence of a strongly acidic cation exchange resin, especially one of the sulfonic acid type. The process may consist in passing the olefin-containing hydrocarbon stream to the sump of the first reactor and the process water to the sump of the last reactor. In the individual reactors, the process water and the hydrocarbon stream are co-current, however, the process water from the last to the first reactor is cascaded countercurrent with the hydrocarbon stream flowing through the reactor system 35 from the first to the last reactor.

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9 DK 174479 B1 At the top of the first reactor, an aqueous alcohol is removed and the alcohol is extracted by distillation.

Depending on reaction conditions, the recovered alcohol has a purity of 99 to 99.9%. In the reaction of iso-5-butenes, the recovered TBA can contain, in addition to low amounts of sec-butyl alcohol (SBA), also very small amounts of dimeric compounds. Such diisobutenes are contained in the residual C 2 gas and not in the TBA recovered from the aqueous phase.

As with the reaction of propylene according to the invention, the process can also be used for the reaction of iso-olefins as a rubbing process in this manner. Here, the isoolefin-containing hydrocarbon stream enters the top of the first reactor. The process water is fed into the top of the last reactor and it is removed on the sump by this reactor and, contrary to the running hydrocarbon stream, it is again put to the top of the next reactor. At each phase phase separation in the individual reactors, the resulting aqueous phase can be recovered and passed into the further 20 reactors. At the sump of the first reactor, the aqueous crude alcohol is finally extracted and it is worked up distillatively. On the sump of the last reactor is removed a residual gas, now containing only a little isobutene. - _

The process according to the invention exhibits, in the reaction of the isoolefin, because of the countercurrent gas flowing, in relation to the known processes which co-stream the process streams through the reactor system consisting of one or more reaction zones, in addition to large olefin. advantages and great selectivity yet another advantage. Since the distribution equilibrium of TBA between the aqueous phase and the hydrocarbon phase under the selected reaction conditions is approx. 1: 3.5, the residual gas in the co-stream method, according to the prior art, contains large proportions of TBA (12 to 30%).

For the separation of TBA from the residual gas, a multistage extraction or distillation must be used. In the process of the invention, the residual gas when leaving the last reactor contains only 0.2 to 2.5% TBA. In a post-wash with the process water or part thereof (for example in a countercurrent reactor), the remaining 5 gas is obtained practically TBA-free (0.1% TBA), The separation of TBA from the process water is carried out in a distillation column, where at the top of the column an azeotropic TBA is obtained with 12% water and the excess process water is removed on the sump of the column. and it is again fed to the last reactor.

In the process of the invention, the disadvantages of the prior art co-current method can be:

Requirements of a high olefin concentration in the feed gas, insufficient olefin turnover, unsatisfactory selectivity and the necessary post-processing measures such as the distillative ether separation, ether cleavage and alcohol extraction are removed.

The advantages of the countercurrent method of the invention can be used not only for new construction or complete replacement of an existing co-current reactor, but also for the conversion of existing co-current reactors to countercurrent reactors, or by supplementing installations with additional countercurrent reactors, in which olefin reactors charged residual gas from the co-current reactors is reacted such that even without the hitherto post-processing measures, a good olefin turnover of 98 to 99% and a good selectivity of 99% are obtained at satisfactory reactor power.

Thus, by using a co-current system, the method of the invention provides advantages with little investment cost.

In principle, the combination of the co-current and counter-current method can be used arbitrarily by using the solid catalyst systems suitable for direct hydration of olefins. It is also possible to combine other catalyst systems in the stream stream with the preferred acidic cation exchange catalyst in the counterstream reactor.

In the drawings, exemplary embodiments of the invention are illustrated. They are described in more detail below. They show:

FIG. 1 shows a process diagram of the sump method according to the invention; FIG. 2 is a process diagram of the rubbing method 10 according to the invention; FIG. 3 is a flowchart of the prior art co-flow process of the prior art.

In a sump method according to FIG. 1 with a reactor system consisting of four reactors Δ, B, C and D, 15 olefin-containing hydrocarbon mixture (C1 or C4 / C5 section) is fed through line 1 to the sump of reactor A. The process water is connected via conduit 2 to the excess water from the sump of column M from conduit 3, and it is conveyed at conduit 4 to the sump of the last reactor D. In the reaction of C part of the process water is passed through extractor E.

The aqueous phase removed in the upper region of reactor D via conduit 5 is fed to the sump of reactor C, and via conduits 6 and 7 it is fed to the reactors B and A present each time, and finally it is conducted via conduit 8 to column M. The hydrocarbon stream removed at the top of reactor A via line 9 is fed to the sump of reactor B, and then to reactors C and D are fed through lines 10 and 11 each time, and then discharged at propylene reaction via line 12 as residual gas. Conveniently, in the reaction of C ^-C ^ isoolefins, the residual gas is washed prior to Extractor E to remove TBA contained therein. For this purpose, the process water from line 4 or a partial stream of the process water from line 4 is used, which is thus fed together with the remaining process water as described.

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12 DK 174479 B1 the sump of reactor D. Thus, the process water and olefin-containing hydrocarbons are passed through the individual reactors in co-flow, but it is passed through the reactor chain in a countercurrent manner.

The process of the invention can also be carried out as a countercurrent rubbing process. An embodiment is illustrated in FIG. 2 on the basis of the same four-reactor reactor system. In this process, an olefin-containing hydrocarbon mixture is passed through line 21 to the top of reactor A, 10 and the process water is passed through line 24 to the top of reactor D, and the process streams are also cascaded countercurrently through the reactor system.

The process water introduced into reactor D is removed as an aqueous phase via line 25 and it is fed to the top of reactor 15 C and then via lines 26 and 27 to the reactors B and A. each time, the hydrocarbon stream in the opposite direction is removed in the opposite direction. the sump area of reactor A, and via lines 29, 30 and 31 it is fed to the following reactors B, C and D, and then it is discharged via line 32 as residual gas which, in the case of conversion of C -isoolefins are washed to remove therein contained TBA in extractor E.

Either in the upper region (sump method) or in the sump (sump method) of reactor A, the aqueous alcohol is removed via line 8 or line 28 and it is directed to column M, where azeotrope is recovered at the top of the column via line 13 or 33 alcohol which, in the preparation of TBA, may additionally contain 12% water, 0.05 to 1% SBA and very small amounts of dimeric compounds. As needed, the azeotropic alcohol can be dried in known manner.

According to a preferred embodiment of the process, the reaction temperature in the four successive reactors is raised so that even with less water loading, an almost complete olefin turnover 35 (99%) and a selectivity of approx. 99%.

13

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DK 174479 B1

In the C ^-isoolefin reaction, the purity of the produced TBA degraded by the formation of SBA can be improved up to 99/9% when under the conditions of the distillative work-up of the aqueous crude TBA at a suitable site 5 of the azeotroping column. SBA is removed as an aqueous relative to TBA-enriched sidestream via line 14 or 34.

The process of the invention has the advantage that without any concentration of the residual gas, any olefin-containing mixture is reacted almost quantitatively, whether 50% or 98%. The process works even under unfavorable temperature conditions (eg at IPA manufacturing temperatures greater than 150 ° C), very selectively because of the countercurrent mode of operation. Contrary to known co-current-irrigation methods, e.g. in the propylene reaction no isopropyl alcohol and diisopropyl ether with the remaining gas from the reactor system.

The following examples serve to illustrate the invention.

20

Examples 1 to 5 In the embodiment of FIG. In the illustrated process, 1 5320 g (water-moistened) is distributed by a strong acid catibn of a commercial process.

replaces four principals. Via line 1, the first reactor A is fed a propane / propylene mixture. Through conduit 4, the last reactor D is fed with process water that passes through the reactors · cascade-like from the last reactor D to the first reactor A. Flow rates, conditions and results are listed in Table I.

30 35

Table I

14 DK 174479 B1

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Examples 12 2 11

Temperature, ° C 150 150 150 158 158 5 Pressure, bar 100 100 100 100 100

Propylene content, mole% 92 92 92 92 58

Propane content, mole% 8 8 8 8 42 Amount (g / h) process 8000 10000 10000 10000 6000 water (stream 4) Amount (g / t) pro 800 800 1000 1000 1000 10 pan / propylene mixture (stream 1) Amount (mol / h) 17.5 17.5 21.9 21.9 13.8 propylene

Conversion (mol%) 96.7 99.6 90.8 99.3 98.3

Selectivity (mol%) for IPA 99.5 99.4 99.0 99.3 98.9

Selectivity 0.5 0.6 1.0 0.7 1.1

(mol%) for IPE

Selectivity (mol%) 0.1 - <0.1 <0.1 <0.1 <0.1 for iso-hexene plus n-propyl alcohol (NPA) Catalyst Capacity 2.4 2.5 2.8 3, 1 1.9 citation, mole IPA / 1 cat. ♦ t DIPE in the remaining 3.4 4.0 8.1 5.9 6.0 gas, g / t (grease) 25

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DK 174479 B1

Examples 6 to 8

The process according to the invention can also be used as a rubbing process. An embodiment is illustrated in FIG. 2nd

In this process, the propane / propylene hydrocarbon stream is fed via line 21 to the top of the first reactor A and the process water is passed through line 24 to the top of the last reactor D. The aqueous alcohol removed from the sump of the first reactor A wire 28, and the 10 is worked up in known manner.

The four reactors are filled with 5320 g of a strong acid cation exchanger. Flow rates, conditions and results are listed in Table II.

15 20 25 30 35 o

Table II

16 DK 174479 B1

Examples 6 2 1

Temperature, ° C 150 150 158 5 Pressure, bar 100 100 100

Propane in C

Amount (g / t) C 2 -carbon- 800 1000 10 QO

hydride (stream 21) Amount (mole / h) propylene 17.5 21.9 13.8

Conversion (mole%) 97.2 92.9 98.5

Selectivity (mol%) 99.4 99.1 98.5 for IPA 15

Selectivity (mol%) 0.6 0.9 1.5

to IPE

Selectivity (mol%) <0.1 <0.1 <0.1 for iso-hexene plus n-propyl alcohol

Catalyst capacity 2.4 2.9 1.9 20 moles IPA / 1 cat. - t DIPE in the residual gas, g / t (grease) 4.1 7.4 8.5 25 30 35 DK 174479 B1

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17

Examples 9 to 11 (Comparative Examples)

The 3 reactors illustrated are used as co-current reactor reactors under the conditions previously described in Examples 6 to 8.

5 Flow rates, conditions and results are given in Table III.

Table III

Examples 10 1 li '11

Temperature, ° C 150 150 158

Pressure, bar 100 100 100

Propylene content,% 92 92 58 Amount of process water, g / t 8000 10000 6000 (stream 44) 15 Amount of propane / propylene, g / t 800 1000 1000 (stream 41) Amount of propylene, mol / t 17.5 21.9 13.8

Conversion, mol% 74 78 59

Selectivity (mol%) for IPA 91.2 91.2 87 20 Selectivity (mol%) for IPE 8 8 12

Selectivity (mol%) 0.8 0.8 1.0 for iso-hexene plus n-propyl alcohol

Catalyst capacity 1.7 2.2 '1.0 mol ΙΡΛ / 1 cat. T 25 DIPE in the residual gas, 60 81 57 g / h (grease) 30 35 18

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DK 174479 B1

Examples 12 to 14 (Comparative Examples)

Similarly, the four reactors are used as co-current sump reactors under the conditions previously described in Examples 6 to 8.

5 Flow rate, conditions and results are given in Table IV.

Table IV

Examples 12 13 14

Temperature, ° C 150 150 158

Pressure, bar 100 100 100

Propylene content,% 92 92 58 15 Amount of process water, g / t 8000 10000 6000 Amount of propane / propylene, g / t 800 1000 1000 Amount of propylene (mol / h) 17.5 21.9 13.8

Conversion, mole% 72 '74 58

Selectivity (mol%) to IPA 92.4 92.4 88

Selectivity (mol%) for IPE 7 7 H

Selectivity (mol%) 0.6 0.6 0.8

to iso-hexanes plus NPA

Catalyst Capacity, 1.7 2.1 I / O

moles of IPA / 1 cat. * t DIPE in the remaining gas, 51 66 52 25 g / h (grease) 30 35 19

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DK 174479 B1

Examples 15 to 20

For the embodiment shown in FIG. 1, in which a total of 2750 g of a commercially available strong acidic cation exchanger is distributed to four reactors, via line 1 an isobutene-containing C ^ hydrocarbon mixture is passed. Through stream 4, process water is passed through the reactors cascade-like from the last to the first reactor. Flow rates, conditions and results are listed in Table V.

10 15 20 25 30 35

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Table V

20 DK 174479 B1

Examples 15 16 17 18 19 20 5 Temperature, ° C 70 92 92 92 92 98

Pressure, bar 16 16 16 16 16 16

Isobutene content, 45 45 45 78.5 17 45 mol% n-butene content, 40 40 40 8 58 40 mol% 10

Butane, mol% 15 15 15 3.5 25 15 Amount (g / h) 8000 8000 4000 4000 4000 4000 process water (stream 4) Amount (g / t) 2000 2000 2000 2000 2000 2000 (wash water) 15 Amount (g / t) 600 600 600 600 600 600 C 2 -carbon hydride (stream 1) Amount of isobutene 4.8 4.8 4.8 8.4 1.8 4.8 (mol / h) 20 Conversion 86 98.5 92 96 90 98 (mol%) isobutene

Conversion 0.09 1.0 0.9 0.3 1.2 1.3 (mol%) n-butene ~

Selectivity 99.9 99.1 99.1 99.7 98.8 98.7

"(Mol%) for TBA

25

Selectivity 0.09 0.9 0.9 0.3 1.2 1.3

(mol%) for SBA

Selectivity 0.015 0.02 0.02 0.025 0.015 0.025 (mol%) dimeric compounds TBA in a re- <0.1 <0.1 <0.1 <0.1 <0.1 <0.1 sterile gas ( flow 12) weight% 35 21

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DK 174479 B1

Examples 21 to 26

Examples 15 to 20 are repeated with the proviso that the amount of strong acid cation exchange is now distributed to three reactors. Conditions, flow rates and results are given in Table VI.

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Table VI

22 DK 174479 B1

Examples 21 2j? 2J3 2_4 2J5 25

Temperature, ° C 70 92 92 92 92 98 5 Pressure, bar 16 16 16 16 16 16

Isobutene content, 45 45 45 78.5 17 45 mol% n-Butene content, 40 40 40 8 58 40 mol%

Butane, mole% 15 15 15 3.5 25 15 10 Amount (g / h) 8000 8000 4000 4000 4000 4000 (stream 4) Amount (g / t) 2000 2000 2000 2000 2000 2000 Washing water Amount (g / t) 600 600 600 600 600 600.

C 1-6 hydrocarbon (stream 1) Amount (mole / h) isobutene 4.8 4.8 4.8 8.4 1.8 4.8

Conversion 76 90 86 88 82 89 (mol%) isobutene

Conversion (mol%) 0.09 0.9 0.8 0.3 1.1 1.1 20 n-Butene

Selectivity 99.9 99.0 99.0 99.6 98.7 98.9

(mol%) for TBA

Selectivity 0.1 1.0 1.0 0.35 1.3 __ 1.1 (mol%) for SBA ~

Selectivity 0.015 0.02 0.02 0.025 0.02 0.02 (mol%) of dimeric compounds TBA in the re- <0.1 <0.1 <0.1 <0.1 <0.1 <0, 1 gaseous gas, weight% (stream 12) 30 35 23

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DK 174479 B1

Examples 27 to 29

Examples 27 to 29 are carried out according to the embodiment of FIG. 2 illustrates a rubbing process in which the four reactors are filled with a total of 2750 g of strongly acidic cation exchanger.

5 Flow rates, conditions and results are listed in Table VII.

10 15 20 25 30 35

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Table VII

24 DK 174479 B1

Examples 27 28 29

Temperature, ° C 70 92 98 5 Pressure, bar 16 16 16

Isobutene content, mol% 45 45 45 n-Butene content, mol% 40 40 40

Butane, mole% 15 15 15 Amount (g / t) of process water 8000 8000 4000 (stream 24) 10 Amount (g / t) of wash water 2000 2000 2000 Amount (g / t) C 2 -carbon-600 600 600 hydride (stream 21 ) Amount (mole / h) isobutene 4.8 4.8 4.8

Conversion (mol%) 87.5 98.7 98.5 isobutene

Conversion (mole%) 0.2 1.0 1.5 n-butene

Selectivity (mole%) 99.8 99.0 98.5

for TBA

Selectivity <mol% 0.2 0.2 1.5

20 for SBA

Selectivity (mol%) 0.03 0.05 0.05 dimer compounds TBA in the residual gas, weight% (stream 32) 0.1 0.1 0.1 25 25 35 25

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DK 174479 B1

Example 30

In one of FIG. 1 illustrates the procedure in which the reactors are filled with a total of 2750 g of a strongly acidic cation exchanger of a commercial process. per hour of a C 1-6 hydrocarbon mixture containing 45 mole% of isobutene and via line 4 8000 g of hourly process water cascade countercurrent through the reactors.

At an execution temperature of 70 ° C and an output pressure of 16 bar, an isobutene conversion of 86% is obtained. Accordingly, via line 8 8232 g / ml is removed. per hour, aqueous reactor starting product containing 3.72 weight percent TBA, and only traces of SBA and dimer products.

This aqueous reaction product is used for continuous distillation of TBA on the 12th bottom of a total of 50- bound column. At an application rate of 3950 g per per hour aqueous product is removed as the main product 348 g per hour. hourly azeo-trop TBA with 12% water and an anhydrous TBA calculated 99.9% by weight. On the sump of the distillation column 20 an alcohol-free process water is recovered at a set reflux / distillate ratio = 2.

Example 31

In one of FIG. 1 illustrates a process in which the reactors are filled with a total of 2750 g of a strongly acidic cation exchanger of a commercial process, correspondingly Example 6 is passed via line 1,600 g per meter per hour of a C 1-6 hydrocarbon mixture containing in 45 mole% of isobutene and via conduit 4 4000 g / h. hour process water J cascade countercurrent through the reactors.

At an execution temperature of 98 ° C and an execution pressure of 16 bar, an isobutene turnover of 98% is obtained. Accordingly, via conduit 8 4270 g / ml is removed. hourly aqueous reactor starting product containing 8.16 g weight percent TBA, 0.1 weight percent SBA and less than 0.01 weight percent dimers.

35 26 DK 174479 B1

ISLAND

This aqueous reaction product is used for continuous distillation of TBA on the 12th bottom of a 70-bonded column. At an application rate of 4270 g per hourly aqueous product is removed on the 17th floor via line 14, 5 bottoms above the product feed, 14.7 g. per hour of a product stream with 29.2 wt% SBA, 49.5 wt% TBA and 21.3 wt% water. At the same time, at the top of the column 387.9 g per per hour as distillate with 11.96 wt% water and a 99.9 wt% anhydrous TBA.

10 At the sump of the column, at a set reflux / distillate ratio = 2, an alcohol-free process water is recovered.

Examples 32 to 34 (Comparative Examples) 15 For the Comparative Examples, the four reactors are used as the co-flow method (Fig. 3). As catalyst, 2750 g of a strong acid cation exchanger is used. In these experiments, the isobutene-containing C4 hydrocarbon mixture via conduit 41 and the process water via conduit 44 are added to the top of reactor A, and aqueous TBA and the refinery on the sump of reactor D are removed via conduit 48 or conduit 49. Reaction conditions, flow rates and results are listed in Table VIII.

25 30 35

Table VIII

ISLAND

27 DK 174479 B1 (Comparative examples)

Example 32 Example 33 Example 34 Temperature, ° C 70 92 98

Pressure, bar 16 16 16

Isobutene content, mole% 45 45 45 n-Butene, mole% 40 40 40

Butane, mole% 15 15 15 Amount (g / h) 10 process water 8000 8000 4000 (stream 44) Amount (g / t) hydrocarbon (stream 41) 600 600 600 Amount (mole / h) isobutene 4.8 4, 8 4.8 15 Conversion (mole%) 54 62 68 isobutene

Conversion (mol%) 0.2 1.5 1.8 n-butene

Selectivity (mole%) 99.6 97.4 97.0

for TBA

Selectivity (mol%) for SBA 0.35 2.3 2.6

Selectivity (mol%) 0.05 0.3 0.4 for dimeric compounds ''.

TBA in the remaining 12.8 24.1 28.9 25 gas, wt% (stream 49) SBA in the remaining <0.1 0.35 0.4 gas, wt% (stream 49) 30

Claims (10)

A process for preparing isopropyl alcohol and tertiary Ci alkohol alcohols by reacting a hydrocarbon stream containing propylene or C ^ is isoolefins with water in the presence of a strong acidic solid. catalyst at elevated pressure and elevated temperature, discharge of the residual gas and aqueous alcohol product and recovery of the alcohol, characterized in that the hydrocarbon is fed into one end and the process water at another end of a reactor chain consisting of several consecutive 'reaction zones, and both process streams are passed through the reactor chain in countercurrent, however, through the individual reactors in co-current, Process according to claim 1, characterized in that the reaction is carried out in 2-10, preferably 3-5, reactors interconnected. Process according to claim 1 or 2, characterized in that the hydrocarbon stream and the process water are fed to the sump of the individual reactors. The method of claim 1 or 2; characterized in that the hydrocarbon stream and process water are fed to the tops of the individual reactors. Process according to one of the claims 1-4 for the preparation of tertiary alcohols, characterized in that the remaining gas is washed free of tertiary alcohol with process water or part of the process water. Process according to claims 1-5 for the preparation of tertiary alcohols, characterized in that the secondary alcohol formed as a by-product is removed as a sidestream by the azeotropic recovery of the tertiary alcohol from the distillation column. Process according to claim 6, characterized in that the lateral flow is carried away few bottoms, preferably 1-7 bottoms, above the initial bottom of the aqueous alcohol product. 35 O DK 174479 B1 Process according to one of the preceding claims, characterized in that the process is carried out in a reactor column with several superimposed reactor zones. Process according to one of the preceding claims, characterized in that, in one or more known reaction zones used in the co-current method, two or more are displaced two or more in the counter-current method according to the preceding claims and the method is first carried out. in co-current and then in counter-current. Process according to one of the preceding claims, characterized in that the reaction is carried out in the presence of a strong acidic cation exchange resin as a solid-15 catalyst. 20 25 30 35