RU2206560C1 - Method for preparing secondary butyl alcohol - Google Patents

Abstract

FIELD: organic chemistry, chemical technology. SUBSTANCE: invention relates to the improved method for preparing secondary butyl alcohol that is an intermediate product for production of methyl ethyl ketone. Method involves liquid-phase alkylation of n-butenes taken as butane-butene fraction with acetic acid to yield sec.-butyl acetate, the following hydrolysis in the presence of immobile layer of catalyst sulfocationite resin in H⁺-form to yield sec.-butyl alcohol and acetic acid and isolation of sec.-butyl alcohol in system of fractionating columns where firstly mixture of sec.-butyl alcohol. Sec. -butyl acetate and water is separated from acetic acid aqueous solution, the following water removal from this mixture and separation of mixture of sec.-butyl alcohol and sec.-butyl acetate to isolate sec.-butyl alcohol. Hydrolysis of sec. -butyl acetate is carried out into vertical hydrolysis reactor with catalyst layer height 3-12 m where sec.-butyl acetate and water are fed by countercurrent with volume feeding rate 0.42-0.51 h^-1 and extraction of formed sec.- butyl alcohol and acetic acid with unreacted sec.-butyl acetate is carried out and water is fed in the amount compensating its consumption for hydrolysis reaction and for its dissolving in an organic extract. Separation of mixture of sec.-butyl alcohol, sec.-butyl acetate and water from an aqueous acetic acid is carried out in column of azeotropic rectification with efficiency 20-30 theoretical plates by such manner to provide temperature on control plate of exhausting part of column maintaining at the level 101.5-103.9 C in feeding heating steam into vat in the amount corresponding to specific consumption 0.507-0.517 kg/kg of raw. The following removal of water from mixture of butyl alcohol, sec.-butyl acetate and water in column of azeotropic drying with efficiency 22-30 theoretical plates is carried out by manner to provide temperature on control plate of concentrating part of column maintaining at the level 96.1-98.8 C in feeding heating steam into vat in the amount corresponding to specific consumption 0.576-0.702 kg/kg of raw with removal of organic phase in the range from 1.4 to 4.5 wt.-% of amount the raw feeding into this column. Further, the prepared dried mixture of sec.-butyl alcohol and sec.-butyl acetate is separated into fractionating column with efficiency 50-60 theoretical plates with isolation of sec.-butyl alcohol and from distillate of azeotropic drying impurities of low- boiling alcohols formed in hydrolysis and isolation of butene dimmers from an aqueous phase and from part of organic phase on additional fractionating column with efficiency 20-28 theoretical plates and with recovery of vat hydrolysis product from this column into reactor. Method provides to prepare sec.-butyl alcohol of high quality, to provide stable work of the hydrolysis reactor and to simplify the operation process. EFFECT: improved preparing method. 6 cl, 10 tbl, 1 dwg, 61 ex

Description

 The invention relates to an improved method for producing secondary butyl alcohol (VBS), which is an intermediate for the production of methyl ethyl ketone (MEK) - a solvent in the production of nitro-lacquers, resins and extractant in the production of lubricating oils and oil paraffins.

A known method of producing VBC from n-butenes, taken as a fraction of hydrocarbons or individual compounds, by liquid-phase alkylation of acetic acid (AC) to obtain sec-butyl acetate (VBA), hydrolyzing it in the presence of a catalyst of sulfocationic resin in H ⁺ form and separating VBC by azeotropic distillation (Pat. France 2447896, class C 07 C 31/12, Pat. Great Britain 2041364, class C 07 C 31/12).

VBS is obtained by hydrolysis of VBA in the presence of sulfocationic resin in H ⁺ form at 60-100 ^° C with stirring. The reaction mixture is distilled to obtain a triple azeotrope of VBS-VBA-water, from which 92% VBS is isolated under vacuum. The output of the IWS on the taken WBA is ~ 58.6%.

 The method is not technological and cannot be used to create industrial technology, because with stirring, the sulfation cation resin is destroyed.

 The raw materials used contain 3.6% isobutene. Partially isobutene interacts with CC to form tert-butyl acetate, which, upon hydrolysis, turns into tert-butyl alcohol (TBS). During alkylation, part of butenes and isobutenes dimer. The patent does not resolve the issue of dimer isolation. The obtained VBS has a low concentration and cannot be used for the production of methyl ethyl ketone, although at present about 100% of the produced VBS is used in industry for the production of MEK.

 To obtain pure WBC, the resulting dimers and TBS must be removed. In the specified method, the conditions for obtaining pure VBS and its quality indicators are not given.

A known method for producing VBS by transesterification of VBA with C ₁ -C ₅ aliphatic alcohols in the presence of metal alcoholates of groups I and IV of the Periodic System. The amount of alcoholate 2-10 wt.% On the taken ether and alcohol.

 The disadvantage of this method is the need to use absolute BWA and alcohol, which significantly increases the cost of obtaining BWA and complicates the scheme due to the need for a significant number of columns and the use of azeotropic agents for drying alcohols. In addition, waste of inactive alcoholates is generated in an amount of 2-10 wt.% Of the amounts of ether and alcohol that require disposal (Aut. USSR Certificate 734182).

A known method of producing alcohols (US Pat. USA 4384148, CL 568-907), which includes:
a) alkylation of olefinic hydrocarbons with an organic acid under alkylation conditions in the alkylation zone to form an organic ester;
b) hydrolysis of the organic ether with water under hydration conditions in the hydration zone with the formation of alcohol and ether hydrolysis products containing the released organic acid;
c) stripping of alcohol and products of hydrolysis of ether from the released organic acid;
g) the separation of alcohol from ether under conditions of separation in the separation zone and the withdrawal of alcohol as the target product of the process;
d) thermal decomposition of the ether obtained in stage (d) in the zone of thermal decomposition at a temperature in the range of 500-750 ^o C and pressure from vacuum to 102 kg / cm ² to obtain olefinic hydrocarbons and alcohol.

 From the examples given in this patent, it can be seen that the process is carried out in autoclaves with stirring. The conversion of olefin to ether (ethyl acetate) is low (~ 20%). The selectivity of hydrolysis of the ether to alcohol is also low (89.2%), along with alcohol, ether is formed.

The additional conversion of ether into alcohol is carried out under very stringent conditions (540-710 ^o C). Data on the production of ITS in the patent are not given.

 Closest to the technical nature of the present invention is a method for producing VSS described in the article M.I. Farberova, A.V. Bondarenko, A.V. Vavilova. Journal of Applied Chemistry, 8, p. 1838-1840, 1980

According to this method, WBC is obtained from n-butenes taken in the form of a butane-butene fraction (WBP) by liquid phase alkylation of the WK to obtain WBA, by separating it from WK, by hydrolysis of WBA in the presence of a fixed catalyst layer - sulfation cation resin in H ⁺ form - with the allocation of VBS from a mixture with VBA, criminal code and water by azeotropic distillation.

Experimentally, hydrolysis is carried out in a static installation at a temperature of 100 ^o With a molar ratio of WBA: water = 1: 2.5.

 Based on laboratory studies, the authors of the article give a scheme for obtaining IWS. WBC is prepared as follows: WBA and water are mixed and fed to a hydrolysis reactor. Reaction products containing approximately 22-24 wt. % of water is separated in a system of three distillation columns. In the first column, azeotropic distillation is carried out to distill VBS and VBA from the criminal code, as a result of which an azeotropic mixture of VBC and VBA with water is obtained in the distillate, and as the bottoms of the criminal code with a concentration of 84 wt.%, Which is returned to the stage of production of VBA. The distillate of the first column is layered, the aqueous phase is recycled for hydrolysis of the BWA, the organic layer of the distillate is fed to the azeotropic drying column, where, when the organic phase is fully returned to the column irrigation, the water dissolved in the organic phase of the distillate of the first column is removed, which is returned to the hydrolysis reactor. The bottom residue of the azeotropic drying column is fed to a distillation column to obtain a commodity VBS, where VBS is distilled off into the distillate, and a mixture of unreacted VBA with VBS is removed from the cube, which is returned to the hydrolysis reactor.

 When water and VBA are fed from above the hydrolysis reactor, due to the low mutual solubility and differences in densities, water and VBA are stratified. After a short period of operation, the aqueous phase accumulates in the lower part of the hydrolysis reactor, and the organic phase accumulates in the upper part. Therefore, the process is unstable with a gradual decrease in the conversion of VBA. When the reagents are fed directly into the hydrolysis reactor, a mixture with a large amount of water with respect to the azeotropic composition enters the azeotropic distillation. Separating this water requires energy.

It is known that in the aqueous phase, cation exchangers desulfurize. The desulfurization of H ⁺ form sulfocationites occurs with the participation of water and is accompanied by the release of equivalent amounts of hydrogen and sulfate ions into the solution, which leads to a decrease in the volumetric capacity of sulfocationite and, as a result, loss of activity (P.E. Tulupov. Stability of ion-exchange materials. M .: Chemistry, p. 44, 1984). The desulfurization rate increases with increasing temperature and an increase in the concentration of H ⁺ ions. (Yon Exch Chichester, p. 440-449, 1984).

At a temperature of ~ 100 ^o C in the presence of water and the ester formed by hydrolysis of the ester of CC, sulfocationionite is partially desulfurized. Sulfuric acid formed during desulfurization is concentrated in the bottom residue of the first azeotropic distillation column - aqueous CC, causing corrosion of the equipment.

 Upon receipt of BWA by alkylation of AC with n-butenes, along with the alkylation reaction, a side reaction of dimerization of n-butenes (up to 5 mol% to converted butenes) occurs (V.M. Obukhov, I.P. Stepanova, A.V. Bondarenko, M. I. Farberov, Petrochemicals, XVIII, 2, pp. 262-267).

 BBP, used as raw material for the production of WBC, contains propene, isobutene, which upon alkylation and subsequent hydrolysis form isopropyl alcohol (IPA) and TBS. The described three-column scheme does not allow to clear the WBC from these compounds. The scheme does not contain a description of the column control system, which makes it possible to obtain high-quality VBS.

 The aim of the invention is to obtain high quality WBS, ensuring stable operation of the hydrolysis reactor, simplifying the control process.

This goal is achieved by the method of obtaining VBS from n-butenes taken in the form of BBP, by liquid-phase alkylation of CC with obtaining VBA, separating it followed by hydrolysis in the presence of a fixed catalyst layer - sulfation cation resin in H ⁺ form in VBC and isolating VBC in the distillation column system comprising an azeotropic distillation column, an azeotropic drying column, and a distillation column to isolate IWS. The proposed method differs from the prototype in that the BWA is hydrolyzed in a vertical hydrolysis reactor with a catalyst layer height of 3-12 m, into which the BWA is fed countercurrently with a space feed rate of 0.42-0.51 h ^-1 water and at the same time extraction of the resulting BWW is carried out and CC with unreacted BWA, while the water is supplied in an amount that replenishes its flow rate for the hydrolysis reaction and its dissolution in the organic extract, and the WBB is separated in a distillation column system with the initial separation of the BWB, WBA and water mixture s from the aqueous CC in the azeotropic distillation column with an efficiency of 20-30 theoretical plates in such a way that the temperature on the control plate of the exhaustive part of the column is maintained at 101.5-103.9 ^o With the supply of heating steam to the cube of the column in an amount corresponding to a specific flow rate of 0.507- 0.517 kg / kg of raw material, followed by removal of water from a mixture of VBS, VBA and water in an azeotropic drying column with an efficiency of 22-30 theoretical plates so that the temperature on the control plate of the concentration part of the column is maintained at 96.1-98.8 ^o With the supply to the cube of the column of heating steam in an amount corresponding to a specific consumption of 0.576-0.702 kg / kg of raw material, with the selection of the organic phase in the range from 1.4 to 4.5 wt. % of the amount of raw material supplied to this column, with the separation of the dried mixture of VBS and VBA in the rectification column with an efficiency of 50-60 theoretical plates and the separation of VBS, with the removal of impurities of boiling alcohols and dimers of butenes from the aqueous phase and part of the organic phase formed from the azeotropic distillate drying on an additional distillation column with an efficiency of 20-28 theoretical plates, with the return to the hydrolysis reactor of the bottom product of this column containing BWA.

The hydrolysis of VBA is carried out in a hydrolysis reactor with a layer height of preferably 6-12 m at a volumetric feed rate of VBA of 0.48-0.51 h ^-1 .

 The azeotropic separation of VBS and VBA from the criminal code is carried out in an azeotropic distillation column with an efficiency of preferably 22-28 theoretical plates.

 Azeotropic drying of the mixture of WBC and WBA is carried out in an azeotropic drying column with an efficiency of preferably 25-28 theoretical plates. The separation of low-boiling alcohols and butene dimers is carried out in a distillation column with an efficiency of 22-28 theoretical plates.

 The distillation columns, in the system of which the WBC is separated, consist of two parts: the lower - the exhaustive (distant) part and the upper - concentration (strengthening) part.

 According to the proposed method, as a reflux in a distillation column for separating low-boiling alcohols and butene dimers, a mixture of a hydrocarbon-rich organic phase and an alcohol-rich aqueous phase is used.

 This last procedure allows the most complete separation of impurities with minimal loss of IWS and IW.

 It should be noted that the supply of WBA reagents and water to the countercurrent hydrolysis reactor with simultaneous extraction of the resulting WBC and CC with unreacted WBA ensures stable operation of the hydrolysis reactor.

 The hydrolysis process is controlled by adjusting the specified ratio of BWA-water, phase separation level, hydrolysis temperature and pressure in the apparatus. The selection of the extract is carried out according to the level in the hydrolysis apparatus, i.e. process control is greatly simplified compared to the prototype.

 In the proposed method, the selection of WBC is carried out in a system of distillation columns, the management of which is greatly simplified by determining the position of the control plates in the columns of azeotropic distillation and azeotropic drying and maintaining certain temperatures on these plates. In addition, the effectiveness of all columns was determined. This allows for stable operation of the columns.

 Thus, the combination of all the proposed essential features of the invention allows to obtain high-quality commodity VBS with the content of the main product of 99.95-99.98 wt.%, To ensure the stable operation of the hydrolysis reactor and to simplify the control process.

 Description of the circuit.

 For a better understanding of the invention, the drawing shows a flow chart for the production of IWB, starting from the stage of hydrolysis of IWB obtained by alkylation of the AC with n-butenes taken as BBP.

The drawing shows the positions of the following devices:
1 - heater
2 - hydrolysis reactor,
3 - phase separator
4 - column azeotropic distillation,
5 - capacitor
6 - phase separator
7 - boiler
8 - column azeotron dehydration,
9 - capacitor
10 - phase separator
11 - boiler
12 - distillation column for the allocation of commodity VBS,
13 is a capacitor
14 - boiler ,.

15 - distillation column for the separation of boiling alcohols and dimers of butenes,
16 is a capacitor
17 - reflux capacity,
18 - pump
19 - boiler.

BWA obtained by liquid-phase alkylation of CC with n-butenes taken in the form of BBP is introduced into hydrolysis reactor 2 with a loaded catalyst — sulfocationic resin in H ⁺ form through heater 1.

 BWA is introduced from the bottom of the hydrolysis reactor into the catalyst bed. Fresh and recycle water is introduced from the top of the hydrolysis reactor to the top of the catalyst bed. From the upper part of the hydrolysis reactor, a hydrolyzate stream is removed containing the extract of VBS and CC in the VBA and dissolved water, as well as impurities of low boiling alcohols and butene dimers. The stream is additionally layered in phase separator 3, the lower (aqueous) phase is returned to the reactor, the upper phase is fed to the azeotropic distillation column 4, in which aqueous azeotropes of VBS, VBA, lower alcohols and butenes dimers are separated from the CC, which are separated into condenser 5 into phase separator 6. The aqueous phase is fed to the top of the azeotropic distillation column 4 in an amount necessary for the formation of azeotropes.

The energy in the azeotropic distillation column 4 is fed through a boiler 7. The amount of energy supplied is regulated depending on the temperature on the control plate of the exhaustive part of the azeotropic distillation column 101.5-103.9 ^o C.

 The bottoms product of the azeotropic distillation column 4, consisting of CC, water and high boiling impurities, is taken to the stage of CC alkylation with n-butenes.

 The organic phase from phase separator 6 is fed to the azeotropic drying column 8, on top of which the BBA, VBS and low boiling point azeotropes are removed, which are condensed in condenser 9 and delaminated in phase separator 10. Most of the organic phase is recycled to the azeotropic drying column 8. A smaller part of the organic phase and the aqueous phase is fed to a distillation column 15 to separate boiling alcohols and butene dimers. Energy is supplied to the azeotropic drying column 8 through a boiler 11.

The amount of energy supplied is regulated depending on the temperature on the control plate of the concentration part of the azeotropic drying column 96.1-98.8 ^o C during the selection of the organic phase in the range from 1.4 to 4.5 wt.% Of the amount of raw material.

 The dried and purified mixture of BBC and BWA from the cube of the azeotropic drying column 8 is fed to a distillation column 12, from above which BWA is withdrawn from the cube 13, and from the BWA cube, which is recycled to the hydrolysis reactor 2. Energy is fed into the distillation column 12 through a boiler 14.

 In the distillation column 15, low-boiling alcohols and dimers of butenes are separated from a part of the organic phase and the aqueous phase of the distillate of the azeotropic drying column 8. Azeotropes of low-boiling products are removed from the distillation column 15, which are condensed in the condenser 16. The column is refluxed by distillate. To prevent separation of the distillate, it is mixed in reflux tank 17 by pump 18. The balance amount of distillate is removed from the discharge line of pump 18. Energy is supplied to distillation column 15 through boiler 19.

 Example 1

 A. Obtaining BWA.

A mixture of UK and BBF with a content of n-butenes of 70 wt.% Is fed into a thermostatic tube reactor, into which the sulfocationic resin in H ⁺ form is loaded. The molar ratio of CC: n-butenes is 1.5: 1. The temperature in the reactor is maintained at 90-100 ^{° C.} , a pressure of 10 kg / cm ² . Isolation of BWA is carried out in a packed column. First, unreacted gases are separated from the mixture of CC and BWA, then BWA and BC are separated by azeotropic distillation using water as an azeotropically. The obtained BWA is introduced into the hydrolysis reactor. The concentration of the obtained WBA is given in table. 1.

 B. Obtaining CHD.

0.73 dm ^{3 of a} molded sulfocationite catalyst in the H ⁺ form of a fraction of 2-3 mm is loaded into a hydrolysis reactor 2 with a diameter of 50 mm and a layer height of 0.36 m. To prevent heat loss to the environment, the hydrolysis reactor 2 is thermostated by supplying coolant to the jacket. WBA through the heater 1 is fed into the lower part of the hydrolysis reactor 2 in the catalyst bed. Water is fed to the top of the catalyst bed. The hydrolyzate is taken from the top of the hydrolysis reactor. The phase separation level of the organic mixture (VBS and CC extract in VBA with dissolved water) and aqueous (water with dissolved CC, VBA, VBS) is maintained above the catalyst layer.

 The extract is fed to the packed column of azeotron distillation 4 in order to separate from the UK water azeotropes VBS, VBA, lower alcohols and butenes dimers. The results are shown in table. 1.

 Examples 2-10.

In the hydrolysis reactor 2 described in example 1, load sulfonation catalyst, changing its volume from 1.46 to 23.6 DM ³ in each experiment. The height of the catalyst layer varies from 0.72 to 12 m

 The hydrolyzate is fed into the packed column of azeotropic distillation 4 in order to separate from the UK water azeotropes VBS, VBA, lower alcohols and butenes dimers.

 Example 11 (comparative).

The hydrolysis reactor 2 is charged with the same amount of catalyst as in example 5, but the reactor has a diameter of 100 mm and a height of the catalyst layer of 0.77 m. Under the same conditions as in example 5, the volumetric feed rate of VBA = 0.48 h ^-1 . Conversion decreased to 38%.

 Example 12 (comparative).

 The same volume of catalyst is loaded into the reactor as in Example 5, but the reactor has a diameter of 32 mm and a catalyst bed height of 14 m. Under the same conditions as in Example 5, a normal counterflow is not achieved.

From the table. 1 and examples 11 and 12 it is seen that with increasing height of the catalyst layer increases the conversion of WBA. The WBA conversion reaches 49-50% with a catalyst bed height of 3-12 m and a WBA feed rate of 0.42-0.51 h ^-1 , it is preferable to work with a catalyst bed height of 6-12 m at a WBA feed rate of 0.48- 0.51 h ^-1 .

 Example 13

The hydrolyzate in an amount of 100 kg / h of the following composition, wt.%:
Water - 9.327
Isopropyl alcohol - 0.157
tert-butyl alcohol - 0.254
Secondary Butyl Alcohol - 27,000
sec-butyl acetate - 44.035
Acetic acid - 18.851
Butene Dimers - 0.376
fed to the azeotropic distillation column 4 with an efficiency of 27 theoretical plates per 6th theoretical plate from the top of the column. As a control plate use the 4th theoretical plate of the exhaustive part of the column (10th theoretical plate, counting from the top of the column). The specific consumption of heating water vapor is maintained in an amount of 0.481 kg / kg of raw material. The temperature on the control plate, the output parameters (total flow of the distillate, the ratio of the organic and aqueous phases in the phase separator 6, the temperature of the upper plate and the bottom of the column), the clarity of separation (selection coefficient of WBS and the content of water and HC in the organic phase of the distillate) are controlled.

 The results are shown in table. 2.

 Examples 14-20.

 They work according to the scheme of example 13 (see drawing) with the same composition of raw materials, changing the amount of energy supplied to the cube of the azeotropic rectification column 4.

 The results of the experiments are given in table. 2.

 From the table. 2 shows that a gradual increase in energy supply constantly increases the selection of distillate. At the same time, first (examples 13-15), the proportion of the organic phase in the distillate increases, and the content of CC in the organic phase decreases, and the selection coefficient of CHD increases. With a further increase in energy supply (Examples 16–20), a complete selection of WSS with the organic phase of the distillate is achieved, its amount and water content in it are fixed. In the distillate, the proportion of the aqueous phase increases, i.e. an increase in the flow of distillate increases its circulation. Since all the components that were planned to be taken with the aqueous phase of the distillate are already selected (example 16), a further increase in the intensity of heat supply and, accordingly, the intensity of circulation of the aqueous phase (examples 17-20) does not give an additional positive effect. On the contrary, with an increase in the intensity of heat supply, the content of CC in the target product of this column — the organic phase of the distillate — increases.

 Therefore, the intensity of energy supply to the boiler of the azeotropic distillation column 4 should be limited and the specific consumption of heating steam should be in the range 0.507-0.517 kg / kg of raw material.

 From the table. Figure 2 also shows that the change in the separation mode and clarity does not practically affect the top temperature of the azeotropic rectification column 4, and the cube temperature increases with an increase in the WBS selection coefficient to 100%, but does not change with a further unproductive increase in the heat supply intensity. Therefore, these parameters cannot be used to control the process.

At the same time, the temperature on the 4th theoretical plate of the exhaustive part of the column (10th theoretical plate from the top of the column) responds to an increase in the heat supply intensity over the entire range, and this point can be used as a control plate to control the process. Maintaining the temperature of the control plate in the range of 101.5-103.9 ^o C allows you to fully extract the target components from the organic phase of the distillate, while avoiding unproductive energy costs and keeping the concentration of CC in the organic phase below 0.02 wt.%.

 Examples 21-24.

 The separation is carried out in an azeotropic distillation column 4 described in Example 13.

For separation serves 100 kg / h of a mixture of hydrolysis reaction products of the following composition, wt.%:
Isopropanol - 0.173
tert-butyl alcohol - 0.280
Secondary butyl alcohol - 29,780
sec-butyl acetate - 48.570
Acetic acid - 20.792
Butene Dimers - 0.414
A different amount of water is added to the extract to obtain an irrigated mixture of different composition, which is used as the raw material of an azeotropic distillation column 4. In all examples, a mixture of azeotropes of alcohols, VBA and butene dimers with water is distilled off, the distillate is separated, the aqueous phase is completely returned to irrigation of the azeotropic column rectification 4, the organic phase is completely selected.

 In all examples, the same specific consumption of heating steam 0.514 kg per 1 kg of the total amount of organic components is used, as a result of which in all examples, regardless of the water content in the feed, the same distillate stream of 103.1 kg / h is obtained.

 The results of the experiments at different water contents in the feed of the azeotropic distillation column 4 are presented in table. 3.

 These examples show that the recommended values of the intensity of energy supply to the boiler of the azeotropic distillation column 4 and the operating temperature on the control plate allow the WBC to be completely removed from the organic phase of the distillate regardless of fluctuations in the water content in the column feed. With fluctuations in the water content, the composition of the distillate does not change; excess water is discharged from the column with bottoms.

 Examples 25-30.

 The azeotropic distillation of the mixture according to example 13 is carried out in columns of different efficiency with the same power supply, the same amount of energy supplied to the cube (specific consumption of heating steam in the boiler of the azeotropic distillation column 4 0.514 kg / kg of raw material) and, accordingly, the same distillate selection 103.1 kg / h .

 When changing the efficiency of the azeotropic distillation column 4, the place of input of the raw material is changed (from 4 to 6 theoretical plates from the top of the column), while trying to maintain a ratio of the efficiency of the concentration and exhaustive parts of the column close to 1: 4. As a control plate, choose a 8-10 plate from the top of the column (or 4 theoretical plate of the exhaustive part of the column).

 The characteristics of the azeotropic distillation column 4 in different experiments and the separation results are shown in table. 4. In all experiments, the target components are completely removed with the organic phase of the distillate.

 As can be seen from the table. 4, the efficiency of the azeotropic distillation column 4 is 20-30 theoretical plates, however, with the same energy consumption, the efficiency of this column affects only the content of CC in the organic phase of the distillate.

 In order to ensure the content of CC in it no more than 0.05 wt.%, It is necessary to have a column with an efficiency of not less than 22 and not more than 28 theoretical plates, which is optimal. A further increase in the efficiency of the column (more than 30 theoretical plates) is not justified, since it will give too little effect.

 Examples 31-38.

To the azeotropic drying column 8 with an efficiency of 25 theoretical plates, 100 kg / h of the organic phase of the distillate of the azeotropic distillation column 4 (an irrigated mixture of VBS-VBA and impurities) of the following composition, wt.% Are fed:
Water - 10.32
Isopropanol - 0.20
tert-butyl alcohol - 0.32
Secondary butyl alcohol - 33.71
sec-butyl acetate - 54.98
Butene Dimers - 0.47
The supply of raw materials lead to the level of the 10th theoretical plate from the top of the column. A mixture of aqueous azeotropes of low-boiling alcohols and butene dimers with an addition of VBS and VBA is distilled off into the distillate, the mixture is separated, the aqueous phase is completely removed for purification in distillation column 15. Part of the organic phase is also removed for purification in distillation column 15, and its main mass is fed for irrigation Azeotropic drying columns 8.

 From the cube of the column of azeotropic dehydration 8, a drained and purified mixture of VBS and VBA is removed.

 The results of the experiments in the selection of the organic phase in an amount of 3 wt.% Of the amount of raw materials are presented in table. 5.

 From the table. 5 it is seen that with insufficient supply of energy to the boiler of the azeotropic drying column 8 (Example 31), it is not possible either to completely remove water from the VBS-VBA mixture or to clear this mixture of impurities of low boiling alcohols and butene dimers. A small increase in the specific steam consumption (Example 32) from 0.407 kg / kg of raw material from 0.465 kg / kg of raw material allows to completely remove water and low boiling alcohols from the VBS-VBA mixture and to reduce the content of butene dimers in the bottom product of the azeotropic drying column 8.

 However, with such a content of butene dimers in the bottoms product, it is impossible to obtain VBS of the required quality.

 A further increase in energy supply (examples 33-38) leads to a further sharp decrease in the residual concentration of butene dimers in the mixture of VBS-VBA to thousandths of a percent, which will allow to obtain VBS of the required quality, but is accompanied by an increase in distillation of VBS with distillate, which will require additional costs for it extraction during further processing of the distillate.

 From the obtained results it follows that the required degree of purification of the VBS-VBA mixture from butene dimers (the content of butene dimers in the mixture is 0.004% and below) is achieved starting from the specific consumption of heating steam of 0.576 kg / kg of raw material. With increasing consumption of heating steam to 0.702 kg / kg, the residual content of butene dimers in the VBS-VBA mixture decreases to 0.001%, and a further decrease in the residual concentration of butene dimers in the bottoms product is ineffective. Thus, the recommended interval of the intensity of heat supply corresponds to the specific consumption of heating steam in the range of 0.576-0.702 kg / kg of raw material (examples 34-38).

The temperature on the control plate in the operating interval varies almost linearly and can be used to control the supply of heating steam. When selecting the organic phase in an amount of 3% of the amount of incoming raw materials, the working range of the specific consumption of heating steam corresponds to the temperature range on the control plate 96.1-98.8 ^o C.

 As a control plate use the 6th theoretical plate of the concentration part of the column (6th theoretical plate from the top of the column).

 Examples 39-44.

 They work according to the scheme of examples 31-38 (see drawing) with the same composition of the raw materials of the azeotropic drying column 8, changing the amount of the organic phase with an almost constant supply of energy to the boiler of the column (specific consumption of heating steam is about 0.645 kg / kg of raw material).

 The results are shown in table. 6.

 From the data table. Figure 6 shows that the required degree of purification of the VBS-VBA mixture from butene dimers (the concentration of butene dimers in the bottoms in the range of 0.002-0.004 wt.%) Is achieved by selecting the organic phase in an amount of 1.4 to 4.5% of the amount of incoming raw materials (0.014-0.045 kg / kg of raw material), while the loss of VSS with distillate ranges from 6.02 to 9.26% of its potential in raw materials. With a deeper cleaning, the losses of IWS noticeably increase.

From the table. Figure 6 shows that the dependence of the temperature on the control plate on the amount of selection of the organic phase is almost linear in the range from 1.4 to 4.5 wt. % of the amount of incoming raw materials (0.014-0.045 kg / kg of raw materials). The temperature on the control plate is 96.1-98.8 ^o C (examples 40-43).

 As a control plate, choose the 6th theoretical plate of the concentration portion of the column (6th theoretical plate from the top of the column).

 Examples 45-49.

 Drying the mixture of VBS and VBA according to the scheme of examples 31-38 (see drawing) is carried out in columns of azeotropic drying 8 of various efficiencies with almost the same selection of the organic phase and the supply of energy to the column. For columns of different efficiency (the number of theoretical plates 20, 22, 25, 28 and 30), the input level of the raw material is changed so that the ratio of the efficiencies of the exhaustive and concentration parts is equal to 1.5: 1 or close to it. In all cases, the 6th theoretical plate of the concentration part of the column (the 6th theoretical plate from the top of the column) is chosen as the control plate.

 The results of the experiments are presented in table. 7.

 From the table. 7 it is seen that the separation and purification of the VBS-VBA mixture in the azeotropic drying column 8 is carried out at an efficiency of 20-30 theoretical plates, however, the required degree of purification of the VBS-VBA mixture from butene dimers is achieved, starting from the efficiency of 25 theoretical plates to 28 theoretical plates. An increase in the number of plates from 28 to 30 gives a very insignificant effect and a further increase in the number of theoretical plates will lead to an unjustified increase in capital expenditures.

 Examples 50-54.

 The distillation of the mixture of the distillate of the azeotropic drying column 8 is carried out in a distillation column 15 to separate boiling alcohols and butene dimers.

In the distillation column 15 serves 10 kg / h of a mixture of distillate columns of azeotropic drying column 8 (a mixture of organic and aqueous phases) of the following composition, wt.%:
Water - 64.07
Isopropanol - 1.22
tert-butyl alcohol - 1.97
Secondary Butyl Syrup - 21.11
sec-butyl acetate - 8.71
Butene Dimers - 2.92
A mixture of azeotropes of low-boiling alcohols and dimers of butenes with water is distilled off into the distillate, and the distillate without separation is divided into 2 parts, the larger part of which is fed to the distillation column 15 for irrigation, and the smaller one is taken for processing to remove impurities from the system.

 Separation is carried out in columns of various efficiencies. Selected 0.95 kg / h of distillate, which corresponds to the selection of 9.5% of the amount supplied to the distillation column of the mixture and about 1.6 times the total amount of impurities contained in it to be removed to the distillate. For each distillation column, the smallest reflux value is determined, starting from which the required separation accuracy is obtained, and the heating steam flow rate required for such separation is compared. The clarity of separation is evaluated, on the one hand, by the coefficients of selection of impurities in the distillate, which should be as large as possible, and, on the other hand, by the coefficients of transition to the distillate of the target products (WBA and VBS), which should be as small as possible.

 The results of the experiments are given in table. 8.

 As can be seen from the table. 8, the efficiency of the distillation column 15 for separating low-boiling alcohols and butene dimers is 20–28 theoretical plates, however, it is preferable to operate at 22–28 theoretical plates, since an increase in the number of plates from 20 to 22 leads to a significant reduction in energy consumption, and this dynamics is maintained at a further increase in the number of plates to 26. An increase in the number of theoretical plates above 28 is not economically feasible.

 Examples 55-61.

 The separation of the mixture of WBC and WBA with obtaining commodity WBC is carried out in distillation column 12 (see drawing).

 In the distillation column 12 serves 100 kg / h purified from water and impurities, a mixture of VBS and VBA, containing 38% alcohol and 62% ether. Separation is carried out in columns of various efficiencies.

 34 kg / h of distillate are selected, which corresponds to a selection of 89.7% of the potential of IWF in the diet.

 For each column, the lowest value of the reflux ratio is determined, starting from which a commodity VBS with a concentration of not less than 99.95 wt.% Is obtained.

 From the table. Figure 9 shows that the separation energy costs when the column efficiency is below 50 theoretical plates are unacceptably large, while increasing the column efficiency from 60 to 65 theoretical plates does not have a significant effect. Thus, the optimal efficiency of the distillation column 12 is the efficiency of 50 to 60 theoretical plates.

 In the table. 10 shows the quality indicators of the obtained commodity VBS.

 As can be seen from the data table. 10, get a marketable product with a high content of VBS (99.95-99.98 wt.%). Tk

Claims (6)

1. The method of producing secondary butyl alcohol from n-butenes taken in the form of a butane-butene fraction by liquid-phase alkylation with acetic acid to obtain sec-butyl acetate, followed by hydrolysis in the presence of a fixed catalyst layer - sulfocationic resin in H ⁺ form, to obtain a secondary butyl alcohol and acetic acid, and the separation of secondary butyl alcohol in a distillation column system in which a mixture of secondary butyl alcohol, sec-butyl acetate and water is first separated from aqueous acetic acid to acid, followed by removal of water from this mixture and separation of the mixture of secondary butyl alcohol and sec-butyl acetate with the separation of secondary butyl alcohol, characterized in that the hydrolysis of sec-butyl acetate is carried out in a vertical hydrolysis reactor with a catalyst layer height of 3-12 m, in which countercurrent sec-butyl acetate is fed with a volumetric flow rate of 0,42-0,51 h ^-1 and water, and extraction was carried out simultaneously formed secondary butyl alcohol and unreacted acetic acid sec-butyl acetate, with water served in an amount that replenishes its flow rate for the hydrolysis reaction and its dissolution in the organic extract, and the mixture of secondary butyl alcohol, sec-butyl acetate and water from aqueous acetic acid is separated in an azeotropic distillation column with an efficiency of 20-30 theoretical plates so that the temperature the control plate to be exhaustive of the column was maintained 101,5-103,9 ^o C when applying to the cube of the column of heating steam in an amount corresponding to the specific consumption of 0,507-0,517 kg / kg feed, followed by removal of the meters of water from a mixture of butyl alcohol, sec-butyl acetate, and water in the column to azeotropic dehydration efficiency of 22-30 theoretical plates so that the temperature on the control plate of the concentration of the column was maintained 96,1-98,8 ^o C when applying to the cube of the column of heating steam in an amount corresponding to a specific consumption of 0.576-0.702 kg / kg of raw materials, with the selection of the organic phase in the range from 1.4 to 4.5 wt.% of the amount of raw materials supplied to this column, the resulting dried mixture of secondary butyl alcohol and sec-butyl acetate shared in rectification column with an efficiency of 50-60 theoretical plates, with the release of secondary butyl alcohol, and impurities of boiling alcohols and butene dimers formed during hydrolysis from the azeotropic drying are removed from the aqueous phase and from a portion of the organic phase on an additional distillation column with an efficiency of 20-28 theoretical plates, s returning to the hydrolysis reactor the bottom product of this column containing sec-butyl acetate. 2. The method according to claim 1, characterized in that the hydrolysis of sec-butyl acetate is carried out in a hydrolysis reactor with a catalyst layer height of 6-12 m at a volumetric flow rate of sec-butyl acetate of 0.48-0.51 h ^-1 .  3. The method according to claim 1, characterized in that the azeotropic separation of secondary butyl alcohol and sec-butyl acetate from acetic acid is carried out in an azeotropic distillation column with an efficiency of 22-28 theoretical plates.  4. The method according to claim 1, characterized in that the azeotropic drying of the mixture of secondary butyl alcohol and sec-butyl acetate is carried out in an azeotropic drying column with an efficiency of 25-28 theoretical plates.  5. The method according to claim 1, characterized in that the separation of low-boiling alcohols and butene dimers is carried out in a distillation column with an efficiency of 22-28 theoretical plates.  6. The method according to claim 1, characterized in that as a phlegm in a distillation column separating low-boiling alcohols and butene dimers, a mixture of a hydrocarbon-rich organic phase and an alcohol-rich aqueous phase is used.

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