JPS626526B2 - - Google Patents

Description

DETAILED DESCRIPTION OF THE INVENTION The present invention relates to a method for effectively separating and recovering isobutene from a C ₄ mixture through a predetermined series of processes using an ion exchange resin catalyst.

The standard method using sulfuric acid (U.S. Patent No. 2388971, Japanese Patent Publication No. 41) is used to separate isobutene.
-6165, Special Publication No. 41-7684), Badger-CFR method (Hydrocarbon Processing Vol42, No.11 P186
(1963)) have been implemented industrially, but as is well known, there are many problems such as dealing with the corrosive nature of sulfuric acid and disposing of waste sulfuric acid. In order to avoid these difficulties, a method for isobutene separation that does not use sulfuric acid has been proposed. According to Japanese Patent Publication No. 47-41882, isobutene in a C ₄ mixture is selectively reacted with methanol using a cation exchange resin as a catalyst to convert other C ₄ into methyl tert-butyl ether.
Separated from the distillate and processed on a solid catalyst supporting metals.
A method for recovering isobutene by performing gas phase decomposition at temperatures above 250°C has been proposed, but there are problems with equipment such as equipment difficulties caused by conducting an endothermic reaction at high temperatures and loss of methanol that should be recycled and reused. I have a point.

U.S. Patent No. 4012456 (Japanese Unexamined Patent Publication No. 51-59802) is a method using a cation exchange resin, but
In this case, isobutene is selectively reacted with water under a cation exchange resin catalyst, separated from other _C4 mixtures as tertiary butanol, and isobutene is subsequently recovered by dehydration and decomposition reaction. . However, when reacting water and isobutene, a method is adopted in which a polar solvent and emulsifier are present that dissolve both water and the C ₄ mixture, so it is unavoidable that some amount of the solvent and emulsifier are altered or lost. It is inevitable that the process itself becomes considerably complicated because the solvent and emulsifier are collected and reused. In addition, in the dehydration and decomposition reaction of tertiary-butanol, the reaction is carried out in the gas phase, so the amount of cation exchange resin used as a catalyst is large, making the equipment bulky. Technical difficulties, such as the inability to use high temperature heat sources due to the low upper limit of the heat resistance of the catalyst and the associated need for very large heat transfer surfaces to supply the heat of endothermic reactions with small temperature differences It is unavoidable that production becomes extremely difficult.

The present inventors researched a method for separating isobutene that solved these conventional methods, and found that the following series of processes were performed: (a) Continuously supplying an isobutene-containing hydrocarbon and water to a first reactor; The reactor is filled with strong acid type cation exchange resin catalyst particles as a fixed bed, and the gaps between the catalyst particles are filled with the liquid isobutene-containing hydrocarbon as a continuous phase,
The average linear velocity of water as it travels over the surface of the catalyst particles
A method of flowing down at a rate of 1.0 m/hr or more, at a temperature of 50 to 150
The isobutene-containing hydrocarbon is continuously brought into contact with water at ℃ to perform a hydration reaction and the reactant is extracted from the reactor; (a) the reactant is phase-separated into an aqueous phase and a hydrocarbon phase; By distilling the hydrocarbon phase in the distillation column (1), unreacted hydrocarbons are recovered from the top of the column, and a stream mainly containing tertiary butanol (TBA) is obtained from the bottom of the column. Both the stream mainly containing TBA from 1) and the aqueous phase of the reactant are fed to the distillation column (2), and (d) the stream from the top of the distillation column (2) is discharged to the outside of the system. In addition, a stream mainly containing water is discharged from the bottom of the distillation column (2), a stream containing TBA is obtained from the top of the distillation column (2), and this stream mainly containing TBA is sent to the second reactor. (e) A cation exchange resin catalyst is housed in the second reactor, and a liquid mixture of TBA and water is maintained at a temperature of 90 to 180 degrees Celsius and a pressure of 1.5 to 15 atm. , the stream containing TBA is continuously supplied into this reactor to perform a dehydration and decomposition reaction of TBA, and (f) the generated gaseous isobutene, unreacted gaseous TBA and water vapor are removed from the upper part of the second reactor. (g) This mixed stream is supplied to a distillation column (3) to recover the product isobutene from the top of the column, and the fraction from the bottom of the column is circulated and supplied to a second reactor; (h) On the other hand, a part of the liquid mixture of unreacted TBA and water contained in the second reactor is continuously extracted from the liquid phase portion of the liquid mixture in the second reactor, and (v) this liquid mixture is is circulated and supplied to the distillation column (2).

We have discovered an excellent method for continuous separation and recovery of isobutene by employing the following method.

The method of the present invention will be explained below with reference to the drawings.

The pressurized liquefied isobutene-containing hydrocarbon containing isobutene is continuously fed to the first reactor 1 from the tube 2 . Isobutene-containing hydrocarbons are hydrocarbon mixtures containing isobutene, and generally are mixtures of butenes and butanes containing 10% or more of isobutene. Industrially, a C ₄ hydrocarbon mixture obtained by thermal cracking, steam cracking, catalytic cracking, etc. of petroleum products, preferably one from which butadiene has been separated and removed, is used. The concentration of isobutene in hydrocarbon mixtures is usually below 80%, and generally between 20 and
50%. Hereinafter, the isobutene-containing hydrocarbon will be abbreviated as a C ₄ mixture.

The first reactor has the function of producing tertiary butanol (hereinafter abbreviated as TBA) by performing a hydration reaction between water and isobutene, which are continuously supplied simultaneously from tube 3, and is a reactor in which a strong acid type cation exchange resin is used. Packed as a fixed bed catalyst.

The strong acid type cation exchange resin referred to here has a skeleton of styrene/divinylbenzene copolymer, and the factors that express its properties include exchange capacity, particle size, specific gravity, degree of crosslinking, surface area, and porosity. is important. Among these, a strong acid type cation exchange resin having a surface area, porosity, exchange capacity, and particle size within a certain range is preferably used.

Regarding surface area, cation exchange resin is 80°C.
The value measured using nitrogen according to the surface area measurement method of BET after drying in vacuum for 6 hours is 0.2-120 m 2 /g, preferably 0.4-120 m ² /g.
100m ² /g. This value is 0.2m ² /
If it is less than 120 m 2 /g, the desired hydration reaction rate will be slow, and if this value is more than 120 m ² /g, problems will arise in the durability and mechanical strength of the resin. Regarding porosity, Prac.Natl.Acad.Sci., Vol 7, 115
The value measured using mercury according to the method described in (1921) is 0.03ml/ml or more, preferably 0.05 to 1.0ml/ml.
ml is particularly effective.

A low porosity has the disadvantage that the rate of the desired hydration reaction is slow. Resins that are commonly referred to as gel-type resins generally fall outside of this range, and resins that have predetermined pores in their resin structure are preferably used in the present invention.

Furthermore, regarding the exchange capacity, its value is
Those used have a particle size of 1.0 meq/g or more, preferably 2.0 to 6 meq/g, a particle size of 0.1 to 5 mm, and a true specific gravity of 1.0 to 1.4. The exchange capacity is expressed in milligram equivalents of sulfonic acid groups per gram of dry catalyst particles, and if it is less than 1.0, the rate of the hydration reaction is reduced.

Moreover, if the particle size is smaller than 0.1 mm, it is difficult to put it into practical use as a packed fixed bed, and if the particle size is larger than 5 mm, the activity will decrease. It is desirable that the divinylbenzene content (degree of crosslinking) in the resin skeleton is in the range of 1 to 15%.

Strong acid type cation exchange resins with these properties are good solvents for monomers when copolymerizing styrene and divinylbenzene in a suspended state, but they have poor solvent ability to swell the polymer, such as It can be produced by performing a polymerization reaction using tertiary amyl alcohol, secondary butyl alcohol, isooctane, etc., and then sulfonating the resulting polymer compound.

The first reactor is usually a vertically cylindrical type having a wire mesh or perforated plate at the bottom, on which the granular catalyst is packed. A substantial portion of the space in the reactor other than that occupied by the catalyst is filled with a C ₄ mixture containing isobutene forming a continuous phase. Water, one of the reactants, is continuously supplied through a disperser at the top of the reactor and flows down the surface of the catalyst particles.

Water flows down due to gravity, and also because it has a higher specific gravity than isobutene, which forms the continuous phase.

Isobutene is dissolved in the water through a thin film of water flow covering the catalyst particles and a catalytic reaction of water and isobutene is achieved on the catalyst surface.

In the case of this catalyst mode of the present invention, the hydration reaction of isobutene proceeds in extremely high yield even if the aqueous phase and the hydrocarbon phase remain heterogeneous, and hardly any by-products such as isobutene polymers are produced.

The reason why such an excellent hydration reaction is achieved is not clear, but it is due to the special contact mode on the catalyst surface and the affinity of isobutene and water to the catalyst surface.
It is thought that the state of adsorption has an excellent effect on the progress of the hydration reaction with high yield.

In the figure, the C ₄ mixture forming a continuous phase moves downward continuously from the top to the bottom of the reactor, but the C ₄ mixture is supplied from the bottom of the reactor and moves upward while forming a continuous phase as a whole. However, a method (not shown) of extracting this from the upper part of the reactor is also adopted in the present invention.

For example, the following method may be employed to effectively start the contact method of the present invention.

That is, the reactor is filled with liquid hydrocarbon (isobutene-containing hydrocarbon or inert hydrocarbons, etc.) in advance, and the C ₄ mixture is supplied from tube 2 and water is started from tube 3, and the temperature is gradually increased. Start the reaction. Furthermore, although the C ₄ mixture and water are supplied through tubes 2 and 3, a method may also be adopted in which, at an early stage, a large amount of water is extracted from tube 5, the reactor is filled with the C ₄ mixture, and the temperature is raised.

In the present invention, the hydration reaction temperature is 50 to 150°C.
Preferably the temperature is 60 to 100°C. If the temperature is lower than 50°C, the reaction rate will be slow and TBA will not be effectively produced. If the temperature is higher than 150°C, isobutene dimer, isobutene trimer, secondary As the production of by-products such as butyl alcohol increases, the deterioration of the catalyst becomes more rapid.

In the present invention, the pressure is not particularly limited, but it is employed at a pressure that allows the isobutene-containing hydrocarbon and water to become liquid at the above reaction temperature, and is usually 2 to 50,000,
Kg/cm ² G, preferably 5 to 40 Kg/cm ² G is employed.

In the present invention, it is preferable that the strongly acidic cation exchange resin used is first thoroughly moistened with water. Wetting with water is usually done by immersing the strong acid type cation exchange resin in a large amount of water at 0 to 100° C. for about 1 minute to 24 hours.

In the present invention, when a strongly acidic cation exchange resin that is sufficiently wetted is used, the production of side reaction products, particularly isobutene polymers, is suppressed and the deterioration of catalyst activity is greatly improved.

When using a strong acid type cation exchange resin that has been sufficiently moistened at the beginning in this way, the dry resin is filled into the reaction vessel, water is first supplied to it, and after contact with the water, water is added. A preferred method is to discharge the isobutene and then fill the reactor with isobutene using the method described above.

In addition, in the present invention, water is transported at an average linear velocity of 1.0 m/hr or more, preferably 1.0 to 30 mm/hr, based on the sky column of the reactor.
It is more preferably supplied at a rate of 1.5 to 20 m/hr. When the average linear velocity is less than 1.0 m/hr, the conversion rate of isobutene to TBA tends to be low, and the polymerization of butenes tends to be accelerated.

If the average linear velocity is higher than 30 m/hr, the conversion rate of isobutene to TBA will similarly be low, and the concentration of TBA in the aqueous solution will decrease, which may cause problems in subsequent operations. Further, in the present invention, the C ₄ mixture is supplied at an average linear velocity based on the sky column in the reactor from 0.2 to 50 m/hr, preferably from 1 to 30 m/hr.

The feed rate to the catalyst is 0.1 to 5.0 hr ^-1 , preferably 0.2 to 5.0 hr -1 in LHSV based on water and C ₄ mixture.
4 hr ⁻¹ , most preferably 0.3 to 1.5 hr ⁻¹ .

The reaction mixture that has undergone the catalytic reaction as described above is extracted from the tube 5 and separated into an aqueous phase and a hydrocarbon phase in a static tank 7, but the generated TBA exists across both phases. In this case, a part of the fluid from tube 5 can also be circulated to the reactor via tube 6, which controls the temperature in the reactor and homogenizes the reaction, thereby achieving a more effective reaction. . The circulating volume at that time is LHSV 0.1 to 30, preferably 0.2
~10 gives good results. If it is too small, the desired temperature control etc. will not be carried out effectively, and if it is too large, the conversion rate of isobutene to TBA will decrease.

In the present invention, the above hydration reaction can be carried out in multiple stages. As an example of this case, two hydration reactors are shown in the figure, and the reaction mixture extracted from tube 5 is introduced into static tank 7, separated into a hydrocarbon phase and an aqueous phase, and the hydrocarbon phase is introduced through tube 2' into a first reactor 1' which is similar to the first reactor. At this time, a hydrocarbon phase is removed between the tube 2' and the first reactor 1'.
Installing equipment to separate or remove some or most of the TBA (not shown) will help increase isobutene recovery from the _C4 mixture. In this case, the separated TBA is preferably supplied to the subsequent distillation column (1) or (2).
In addition, when performing the hydration reaction in a multi-stage hydration reactor,
It is preferable to set the reaction temperature in the second and subsequent reactors to be equal to or 1 to 20° C. lower than that in the first reactor to obtain a high yield of isobutene. The reaction mixture extracted from the final hydration reactor is separated into an aqueous phase and a hydrocarbon phase in a static tank 7', and the hydrocarbon phase is fed to the distillation column (1) through a pipe 8. . In the distillation column (1), a stream mainly containing TBA comes from the bottom of the column, and an unreacted stream from which most of isobutene has been removed comes from pipe 12 from the top of the column.
A C ₄ mixture is obtained, which can be effectively used as a raw material for chemical synthesis. here,
Streams containing mainly TBA are usually 50-95wt% TBA,
Unreacted hydrocarbons 0.01-3.0wt%, water 5-50wt%
This includes: The stream mainly containing TBA from the bottom of the column is fed to the distillation column (2) through pipe 13 together with the aqueous phase containing TBA extracted from static tank 7. The function of the distillation column (2) is to remove physically dissolved butane and butenes, maintain the quality of the final product isobutene at a high level of purity, and maintain the concentration used in the next step of dehydrating and decomposing TBA to recover isobutene. It has three functions: concentrating TBA to a maximum of 100%, and collecting hot water from which TBA has been removed and recycling it to the first reactor for thermal economy.

In addition, the present invention simultaneously processes three streams: the aqueous phase stream from the static tank, the bottom stream from the distillation column (1), and the stream from the lower part of the second reactor (pipe 21), which will be described later. This is a major feature. As mentioned above, the bottom stream of the distillation column (1) contains a small amount of unreacted hydrocarbons, which adversely affects the purity of the final product, isobutene. Of course, it is not impossible to perform sufficient distillation in the distillation column (1) so that almost no unreacted hydrocarbons are contained in the bottom stream of the distillation column (1), but in this case, the distillation column (1) ), which causes a large amount of TBA to flow out into the tower overhead stream, which is extremely disadvantageous economically. In the present invention, the bottom stream of the distillation column (1) containing a small amount of unreacted hydrocarbons is fed to the distillation column (2) simultaneously with the aqueous phase from the static tank and the flow from the second reactor as described above. One of the characteristics is that end-reacted hydrocarbons are very effectively removed from the top of the distillation column (2) by distillation.

Physically dissolved butane and butenes are removed from the top of the distillation column (2), and a TBA-water azeotrope or a TBA aqueous solution with a TBA concentration of 50 to 88 wt% is obtained as a side cut from the top of the column. It is convenient. The stream supplied from pipe 13 to the distillation column (2) also contains a small amount of secondary butanol (hereinafter abbreviated as SBA) produced as a by-product in the first reactor, but the concentrated
If the SBA content in the TBA aqueous solution is suppressed to 20 wt% or less, isobutene of sufficiently high purity can be recovered in the subsequent process of the present invention, and the range of allowable SBA content in the concentrated TBA aqueous solution is sufficiently wide. It is possible to take

The concentrated TBA aqueous solution from the distillation column (2) is normally heated through a pipe 16 and then transferred to the second reactor 1.
0 (dehydrolysis reactor), preferably in its lower part. A stream mainly containing water is withdrawn from the bottom of the distillation column (2), and although this stream can be discharged outside the system, it is circulated to the first reactor to collect the water used for the hydration reaction. It can be used as part or all of.

The dehydration decomposition reactor used in the present invention includes at least
Raw material TBA supply port, product and unreacted at the top
It is a pressure-resistant sealed container with an outlet for the TBA mixed gas and an outlet for the liquid mixture of TBA and water at the bottom.

One of the features of the present invention is that the dehydration decomposition reactor can be equipped with no equipment for supplying or removing heat by selecting the reaction conditions as described later; There is no hindrance to having equipment. The reactor is filled with a cation exchange resin catalyst, and a liquid mixture containing TBA and water is housed in contact with the catalyst. This liquid mixture is preferably present as a liquid phase in most of the reactor, particularly in the area where the catalyst is present. This liquid mixture usually contains about 2 to 70 wt% TBA, preferably 10 to 60 wt%.
It is an aqueous solution containing %. The catalyst may be present as a fixed bed or in a suspended or fluidized state in a liquid mixture of TBA and water.

Providing a stirring blade in the reactor or circulating the liquid mixture with a pump to maintain the catalyst in the reactor in a suspended and fluidized state gives favorable results to the dehydration and decomposition reaction of TBA. This is particularly convenient for continuously replacing deteriorated catalysts with new catalysts.

A cation exchange resin is used as a catalyst for dehydration and decomposition, and the cation exchange resin referred to here is a resin that has an acid group and has cation exchange performance, such as a styrene sulfonic acid type resin or a phenolsulfonic acid type resin. Strong acid type resins such as , etc. are representative of these, and can be used if the hydrogen ion exchange capacity is 0.1mM/g or more. As one method that meets this purpose, it is also possible to use the waste catalyst from the first reactor (hydration reactor) as it is as a catalyst in the dehydration cracking reactor, and the method for using such a catalyst is as follows. This has a great effect on increasing the economic efficiency of this process as a whole.

The temperature inside the dehydration decomposition reactor containing the catalyst and the liquid mixture is 90 to 180°C, preferably 105 to 180°C.
The temperature is maintained at 140°C and a pressure of 1.5 to 15 atmospheres, preferably 3 to 10 atmospheres.

As mentioned above, the raw material TBA is heated and supplied into the reactor in a liquid or gaseous state, but it is supplied at a pressure higher than the internal pressure of the reactor, for example, about 0.1 to 5 atm, and the temperature is lower than the internal temperature of the reactor. It is also preferable to supply at a high temperature, for example, about 1 to 50°C. However, in order to control the reaction, feeding at a temperature about 1 to 30° C. lower than the internal temperature of the reactor is sometimes adopted. When supplying TBA in gaseous form
TBA is introduced into the liquid mixture in the reactor, and is usually 1/1/2 of the height of the liquid phase of the liquid mixture.
It is introduced in 3 or less places, preferably 1/2 to 9/10 places. If gaseous TBA is introduced into the upper part of the reactor, the supplied gaseous TBA will not react sufficiently and the amount of unreacted TBA drawn out from the upper part of the reactor will increase.

The introduced TBA comes into contact with the liquid mixture of TBA and water that already exists at the predetermined temperature and pressure, is absorbed and condensed therein, and simultaneously undergoes a dehydration and decomposition reaction on the surface of the catalyst.

If the temperature in the reactor is too low, the amount of isobutene produced per unit catalyst will decrease, and if it is too high, catalyst deterioration will be accelerated and there will be disadvantages in using a high temperature heat source. Furthermore, if the pressure is too low, the raw material TBA supplied in the form of a gas will not be sufficiently absorbed and condensed into the mixed liquid in the reactor after being supplied, and an effective reaction will not take place. If the pressure is too high, the partial pressure of isobutene produced in the reaction system will increase, making it difficult to carry out the dehydration and decomposition reaction in an equilibrium manner.

The isobutene produced by the reaction is continuously withdrawn in gaseous form from the top of the reactor via pipe 18. This gaseous isobutene contains unreacted gaseous TBA and also contains a small amount of water produced by the decomposition reaction in the form of water vapor. Furthermore, a trace amount of a low-amount polymer (mainly a dimer) of isobutene produced as a by-product from the reaction may be contained.

In the present invention, the conversion rate of TBA to isobutene is operated at 20 to 95%, preferably 30 to 90%, and more preferably 40 to 80%, but if the conversion rate in one reactor is not large. Using two or three reactors in series is not only effective in increasing the conversion rate, but also helps to prevent a decrease in production when replacing the catalyst.

The mixed gas extracted from the top of the reactor usually contains 1/10 to 2 times the weight of isobutene, preferably 1/1
TBA is contained in an amount of 5 to 1 times the weight, but if the amount of TBA is excessive, it can be controlled by raising the reaction temperature, moving the raw material TBA feed port to a lower position, etc.

Product isobutene is recovered from a stream containing gaseous isobutene withdrawn from the top of the reactor. This recovery is accomplished by conventional distillation. That is, this stream is introduced into the distillation column (3), and highly pure isobutene can be obtained from the top of the column through tube 20, and a stream containing a small amount of liquid TBA and other impurities can be obtained from the bottom of the column through tube 25. This stream can be heated and fed to a dehydration cracking reactor, but preferably this stream is fed to a distillation column (4) and passed from the top of the column through pipe 27 to remove most of the isobutylene dimer. Most of the bottom of the column can be partially bled through the pipe 28, heated while removing SBA, and recycled to the dehydration cracking reactor.

On the other hand, a portion of the liquid mixture containing TBA and water present in the reactor is continuously extracted from the lower part of the dehydration decomposition reactor through the pipe 21. This withdrawal allows the level of the liquid phase in the reactor to be kept constant while the reaction can be effectively carried out continuously.

The flow extracted from this is the unreacted
TBA, water, and most of the impurities contained in raw materials. When the reaction is operated continuously and at a steady state, the composition of the mixture of TBA and water in the reactor is similar, and a large portion of the water produced by the reaction and unreacted TBA are It is extracted by the flow. This stream 21 can normally be recycled to the reaction. That is, this stream 21 is fed to the distillation column (2).

In the present invention, isobutene is regenerated from TBA produced by the hydration reaction by the method described above, but in the case of the present invention,
In particular, it does not require high temperature or high pressure, produces isobutene in high yield with few by-products, and there is no or little need to supply the heat necessary for the reaction with external heat, so it is carried out using a large reactor. This is possible and industrially advantageous. Further, in the present invention, unreacted TBA is discharged from two places, but since both can be recovered and recycled with a simple operation, no inconvenience occurs.

Furthermore, in the method for producing isobutene according to the present invention, even if the raw material TBA contains impurities such as secondary butanol, it is extremely unlikely that this alcohol will be decomposed to produce n-butenes. High purity isobutene can be produced.

Furthermore, the characteristics of this process include by-products,
All of the contaminants physically dissolved in the _C4 fraction are insoluble or poorly soluble in water, making them easy to remove, making it easy to maintain a high level of purity in the isobutene product, and reducing the amount of contaminants contained in wastewater. It is possible to suppress the content of organic substances to an extremely small amount using normal methods, and has the advantage that the cost of environmental measures is low. In this process, the highest temperature part
Since the temperature is kept below 180℃ and the pH level is kept above 3.5, there will be no corrosion problem and the construction cost of the equipment will be reduced. The ability to use it in a reactor (dehydrolysis reaction) significantly increases the economics of the process.

The present invention will be explained in more detail with reference to Examples below.

Example 1 A cylindrical hydration reactor with a diameter of 14 cm and an internal volume of 60 ml was filled with a fixed bed of 50 ml of a strongly acidic cation exchange resin pre-wetted with a large amount of water as a catalyst. The strong acid type cation exchange resin used here is a suspension polymerization of styrene divinylbenzene using tertiary amyl alcohol as a solvent, and has an exchange capacity.
It has a particle size of 3.6 meq/g, a particle size of 0.4 to 0.6 mm, a degree of crosslinking of 14%, a surface area of 3.4 m ² /g, and a porosity of 0.11 ml/ml. The reactor is continuously fed with a liquid C ₄ mixture containing isobutene and water using a metering pump. The C ₄ mixture used was one obtained by extracting butadiene from the fraction obtained by steam cracking of naphtha, and its composition was as follows.

Isobutane 5.0wt% N-Butane 14.0〃 Trans-butene-2 7.2〃 Isobutene 43.1〃 Butene-1 25.4〃 Cis-butene-2 5.2〃 Butadiene 0.1〃 Prior to the start of the reaction, water and C ₄ mixture were added to tube 12 from tube 3 and tube 2, respectively. /hr, heated at 25/hr and supplied through tube 6 from the reactor outlet to 100%
The temperature inside the reactor was increased to 85% while circulating the fluid at 85%/hr.
℃ to form a continuous phase of _C4 mixture. When the predetermined temperature is reached, the desired reaction proceeds at the above flow rate, and the pressure inside the hydration reactor is maintained at a gauge pressure of 30 Kg/cm ² by continuously withdrawing the amount of liquid corresponding to the amount supplied. do. The linear velocity of water at this time is 2.9 m/hr based on the sky tower. The amount of liquid from the reactor outlet matches the amount supplied from tubes 2 and 3, and is sent to a cylindrical static tank 8 where it is separated into a hydrocarbon phase and an aqueous phase. The hydrocarbon phase is passed through tube 8 and distilled at a reflux ratio of 1 in a distillation column (1) with 20 theoretical plates, with the raffinate C ₄ mixture coming from the top of the column and the mixture from the bottom of the column.
Collect TBA and water. Roughinate _C4 mixture was recovered at a flow rate of 16.8/hr, and the isobutene content therein was 19.1 wt%. The flow from the bottom of the column is supplied through pipe 13 to a distillation column (2) with 30 theoretical plates together with the aqueous phase in the still tank, where distillation is carried out at a reflux ratio of 2.5, and a small amount of _C4 component and isobutene dimer are removed from the top of the column. The hot water containing about 200 PPM of SBA is recovered from the bottom of the tower. In addition, a mixture of TBA-water-SBA with a near azeotropic composition was introduced from the 10th stage from the top of the column.
7.02Kg/hr (TBA: 83.7wt%, SBA: 3.8wt%,
Water: 12.5 wt%) is extracted and converted into a mixed vapor at 135°C in an evaporator via pipe 16, and continuously supplied to the lower part of the dehydration and decomposition reactor through a disperser. The dehydration decomposition reactor is a cylindrical type with a diameter of 16 cm and a capacity of 10 cm, equipped with a stirrer, and inside contains 500 g of a strong acid type cation exchange resin of the same type as that filled in the hydration reactor.
TBA27.7wt% aqueous solution (contains 1.4wt% SBA)
are present as a solid-liquid mixed phase in a total volume of 8.
At steady state, the temperature of the dehydration cracking reactor is 119
℃, and the pressure was 5 Kg/cm ² gauge pressure. A tube (tube 18) is installed at the top of the reactor to extract produced isobutene, unreacted TBA, water vapor, by-produced isobutene dimer, etc. as a gas, and the pressure inside the reactor is maintained at 5 kg/cm ² G. Equipped with an automatic pressure regulating valve.

In addition, the pipe (pipe 21) that removes the mixed liquid portion of TBA, water, and SBA from the catalyst through a screen from the bottom of the reactor has an automatic regulating valve to maintain the level of the solid-liquid mixture in the reactor at 8. is provided. The gas mixture obtained from the top of the dehydration decomposition reactor was distilled at a reflux ratio of 1.5 in a distillation column (3) with 30 theoretical plates to recover 4.42 kg/hr of isobutene. The impurities contained in this isobutene were 110 PPM of butanes, 130 PPM of butene-1, 2250 PPM of butene, 15 PPM of isobutene dimer, and a small amount of water, but the purity of isobutene based on standards excluding water was 99.95. It was wt%.
The bottom fraction of the distillation column (3) is sent through a pipe 25 to a distillation column (4) having 10 theoretical plates, and 0.019 Kg/hr of isobutene dimer is removed from the top of the column at a reflux ratio of 1. Distillate from the bottom of the tower (TBA84.3wt% aqueous solution)
The SBA is partially bled to remove SBA, vaporized in an evaporator from pipe 28, and recycled together with the flow from pipe 16 to the dehydration and decomposition reactor.

The liquid phase portion withdrawn from the dehydration cracking reactor via pipe 21 contains 27.7 wt% TBA and is distilled together with the flow from pipe 13 in the distillation column (2).
An aqueous solution containing 87.3 wt% of TBA is recovered, heated through pipe 16, and circulated as steam to the dehydration decomposition reactor. Hot water was recovered from the bottom of the distillation column (2), but most of it was circulated and supplied to the hydration reactor as a water supply source.

As a result of continuously maintaining the above conditions, the recovery rate of isobutene from the C ₄ mixture was 68.4%, and the selectivity was 99.5%. Furthermore, no deterioration was observed in the cation exchange resin used as a catalyst even after 2000 hours.

Example 2 Hydration reactor 3 with the same shape as shown in Example 1
Each base was filled with 50 strong acid type cation exchange resins as a fixed bed. The strong acid type cation exchange resin used here is obtained by suspension polymerizing styrene divinylbenzene using tertiary amyl alcohol as a solvent, and has an exchange capacity of 3.2 meq/g, a particle size of 0.2 to 0.8 mm, a degree of crosslinking of 10%, and a surface area of 0.72. m ² /g, porosity 0.06ml/ml
has. Prior to the start of the reaction, the reactor is cooled to 5° C. and the C ₄ mixture containing liquid isobutene is filled into the reactor to form a continuous phase of the C ₄ mixture.
Thereafter, the _C4 mixture and water were gradually increased to the top of the first hydration reactor through tubes 2 and 3 until finally 25
/hr, 12/hr. At this time, the temperature of the first hydration reactor is maintained at 90° C. and the pressure is maintained at 20 Kg/cm ² G. The C ₄ mixture used here has the same properties as used in Example 1. through tube 6
While circulating the fluid at a rate of 100/hr, the amount of fluid commensurate with the supply amount is sent to a cylindrical static tank, where it is separated into a hydrocarbon phase and an aqueous phase. In this state, the linear velocity of water in the reactor based on the sky column is 2.9 m/hr. The hydrocarbon phase extracted through pipe 2' is subjected to a flashing operation to remove TBA, and the removed TBA is sent to the distillation column (2), and the C ₄ mixture is cooled to become a liquid phase again.
The second hydration reactor is fed through tube 2', and at the same time 12/hr of heated water is fed through tube 3'.
A continuous liquid hydrocarbon phase has already been formed in the second hydration reactor in the same manner as in the first reactor, and the reactor temperature is maintained at 90°C and pressure at 20Kg/cm ² G. has been done. The hydration reaction proceeds in the same manner as in the first hydration reactor, and the fluid extracted from the second hydration reactor is again separated into a hydrocarbon phase and an aqueous phase in a second static tank. Pipe 2'' (not shown) is directly inserted into the third hydration reactor.
As in the case of the second hydration reactor, 12/hr of water is fed through tube 3'' (not shown). In the third hydration reactor, a continuous phase of liquefied hydrocarbons has already formed in the same way. The temperature, pressure, and amount of circulating fluid are also
The same applies to the second hydration reactor.

The hydrocarbon phase from the third stationary tank is passed through pipe 8 and distilled at a reflux ratio of 1 in a distillation column (1) with 20 theoretical plates. Roughinate C ₄ mixture is removed from the top of the column, and TBA and water are removed from the bottom of the column. to recover. Roughinate C ₄
The mixture was recovered at a flow rate of 14.4/hr, and the isobutene content therein was 3.5 wt%. The flow from the bottom of the column is supplied through pipe 13 to a distillation column (2) with 30 theoretical plates together with the water phase in the three still tanks and the flash residue of the hydrocarbon phase in the first still tank, and the reflux ratio is 2.0.
A small amount of _C4 component, isobutene dimer, etc. are removed from the top of the column, and about 510 PPM is distilled from the bottom of the column.
Collect hot water containing SBA. In addition, from the 10th stage position from the top of the column, a mixture of TBA-water-SBA with a close azeotropic composition was pumped at 9.58Kg/hr (TBA84.9wt%, SBA3.5wt%).
%, water 11.6 wt%) is turned into a mixed vapor at 133° C. in an evaporator via a withdrawal pipe 16, and is continuously supplied to the lower part of the dehydration decomposition reactor through a disperser. The dehydration decomposition reactor has the same structure as in Example 1, and inside the cylindrical shape, 760 g of a strong acid type cation exchange resin used for 8000 hours in the hydration reactor was contained in an aqueous solution of 30.8 wt% TBA (1.1 wt% SBA). ), and the amount of liquid extracted is controlled so that the liquid phase containing the cation exchange resin is 8. In steady state, the temperature inside the dehydration decomposition reactor is 115°C and the pressure is 6 kg/cm ² gauge pressure. At the top of the reactor, there is a tube (tube 18) for extracting produced isobutene, unreacted TBA, water vapor, by-produced isobutene dimer, etc. as a gas.
is installed to control the pressure inside the reactor to 6Kg/cm ² G.
Equipped with an automatic pressure regulating valve to maintain the pressure.

In addition, the pipe (pipe 21) that takes out the mixed liquid portion of TBA, water, and SBA from the bottom of the reactor by removing the catalyst through a screen has an automatic control valve to maintain the level of the liquid phase in the reactor at 8. is provided. The gas mixture obtained from the top of the dehydration decomposition reactor is passed through a distillation column (3) with 30 theoretical plates to a reflux ratio.
1.5 distillations were performed and 6.11 Kg/hr of isobutene was recovered. The impurities contained in this isobutene were 220 PPM of butanes, 160 PPM of butene-1, 210 PPM of butene-2, 65 PPM of isobutene dimer, and a small amount of water, but the purity of isobutene based on standards excluding water was It was 99.93wt%. The bottom fraction of the distillation column (3) is sent through a pipe 25 to a distillation column (4) having 10 theoretical plates, where 0.031 kg/hr of isobutene dimer is removed from the top of the column at a reflux ratio of 1. The bottom fraction (81.1 wt % TBA aqueous solution) is partially bled to remove SBA, and is circulated and supplied to the dehydration and decomposition reactor through pipe 28.

The liquid phase portion withdrawn from the dehydration cracking reactor via pipe 21 contains 30.8 wt% TBA and is distilled together with the flow from pipe 13 in the distillation column (2).
The TBA aqueous solution is collected and circulated through pipe 16 to the dehydration and decomposition reactor. Hot water was recovered from the bottom of the distillation column (2), but most of it was circulated and supplied to the hydration reactor as a water supply source.

As a result of continuously maintaining the above conditions, the recovery rate of isobutene from the C ₄ mixture was 94.5%, and the selectivity was 99.5%.

[Brief explanation of the drawing]

The figure is a flow diagram showing an example of implementing the method of the present invention. 1...First reactor (first stage), 1'...First reactor (second stage), 10...Second reactor, (1)...
... Distillation column (1), (2) ... Distillation column (2), (3)...
... Distillation column (3), 7 ... Still tank (first stage), 7'
...Stationary tank (second stage).

Claims (1)

[Claims] 1. A method for separating and recovering isobutene from an isobutene-containing hydrocarbon mixture, including (a) continuously supplying an isobutene-containing hydrocarbon and water to a first reactor, and containing a strong acid in the reactor. type cation exchange resin catalyst particles are filled as a fixed bed, and the gaps between the catalyst particles are filled with the liquid isobutene-containing hydrocarbon as a continuous phase,
The average linear velocity of water as it travels over the surface of the catalyst particles
A method of flowing down at a rate of 1.0 m/hr or more, at a temperature of 50 to 150
The isobutene-containing hydrocarbon is continuously brought into contact with water at ℃ to perform a hydration reaction and the reactant is extracted from the reactor; (a) the reactant is phase-separated into an aqueous phase and a hydrocarbon phase; By distilling the hydrocarbon phase in the distillation column (1), unreacted hydrocarbons are recovered from the top of the column, and a stream mainly containing tertiary butanol (TBA) is obtained from the bottom of the column. Both the stream mainly containing TBA from 1) and the aqueous phase of the reactant are fed to the distillation column (2), and (d) the stream from the top of the distillation column (2) is discharged to the outside of the system. In addition, a stream mainly containing water is discharged from the bottom of the distillation column (2), a stream containing TBA is obtained from the top of the distillation column (2), and this stream mainly containing TBA is subjected to a second reaction. (e) The second reactor contains a cation exchange resin catalyst, and a liquid mixture of TBA and water is maintained at a temperature of 90 to 180 degrees Celsius and a pressure of 1.5 to 15 atm. The stream containing TBA is continuously supplied into this reactor to perform a dehydration and decomposition reaction of TBA. (g) This mixed stream is fed to the distillation column (3) to recover the product isobutene from the top of the column, and the fraction from the bottom of the column is circulated and fed to the second reactor. , (h) On the other hand, a part of the liquid mixture of unreacted TBA and water contained in the second reactor is continuously extracted from the liquid phase portion of the liquid mixture in the second reactor, and (v) this liquid A method for continuous separation and recovery of isobutene, characterized in that a mixture is circulated and supplied to a distillation column (2). 2. Process according to claim 1, characterized in that a stream mainly containing water from the bottom of the distillation column (2) is recycled to the first reactor. 3. When circulating and supplying the fraction from the bottom of the distillation column (3) to the distillation column (2), this fraction is supplied to the distillation column (4) in advance, and by-product isobutene dimer is discharged from the top of the column. 3. Process according to claim 1, characterized in that the stream from the bottom of the column is recycled to the distillation column (2). 4. Reducing the amount of by-product secondary butanol that accumulates in the system by discharging a part of the flow from the bottom of the distillation column (4) to the outside of the system and circulating the remainder to the distillation column (2). The method according to claim 3, characterized in that: