JP2008162958A - Method for producing highly pure terephthalic acid - Google Patents

Abstract

<P>PROBLEM TO BE SOLVED: To provide a method for producing a highly pure terephthalic acid producing PTA in a good efficiency by forming a highly concentrated crude aqueous terephthalic acid solution by preventing the formation of scale of the terephthalic acid and occlusion by the growth of the same in a heat-exchanger and performing a hydrogenation reaction in a good efficiency. <P>SOLUTION: This method for producing a highly pure terephthalic acid comprising (I) a process of forming crude terephthalic acid slurry by mixing the crude terephthalic acid obtained by the liquid phase oxidation of p-xylene, (II) a process of forming an aqueous solution of the crude terephthalic acid by dissolving the crude terephthalic acid slurry by gradually heating with a multiple number of heat-exchangers, (III) a process of hydrogenation-treating the crude aqueous terephthalic acid solution, (IV) a process of decreasing pressure and cooling the aqueous terephthalic acid solution after the hydrogenation treatment gradually in a multiple number of crystallizing vessels to crystallize terephthalic acid and (V) a solid-liquid separating process of the obtained terephthalic acid slurry is provided by setting ≥1.7 m/sec flow velocity in the heat-exchanger of the crude terephthalic acid at ≤200°C temperature in the process (II). <P>COPYRIGHT: (C)2008,JPO&INPIT

Description

  The present invention relates to a method for producing high-purity terephthalic acid, and more specifically, when preparing a crude terephthalic acid aqueous solution by heating and dissolving a crude slurry of terephthalic acid stepwise in a plurality of heat exchangers. The present invention relates to a method for producing high-purity terephthalic acid, which can prevent clogging due to the generation of scale of terephthalic acid and the growth thereof.

  When paraxylene is oxidized in the liquid phase with a molecular oxygen-containing gas in the presence of an oxidation catalyst such as a catalyst containing cobalt, manganese and bromine in an acetic acid solvent under pressure, terephthalic acid (hereinafter also referred to as “TA”). In addition, crude terephthalic acid (hereinafter also referred to as “CTA”) containing 4-carboxybenzaldehyde (hereinafter also referred to as “4-CBA”) as a main impurity is produced. However, since high-purity terephthalic acid (hereinafter also referred to as “PTA”) is required as a raw material for the production of polyester fibers and the like, it is necessary to purify the CTA.

  As a purification method of CTA, a slurry in which CTA is suspended in water is heated to form an aqueous solution, hydrogenated in the presence of a hydrogenation catalyst, and the treatment liquid is crystallized and solid-liquid separated to produce PTA. The method is known. In general, industrially, in order to sufficiently dissolve CTA in water, it is often carried out stepwise using a plurality of heat exchangers, and the hydrogenation treatment is performed by adding a CTA aqueous solution to the catalyst packed bed. In many cases, the liquid is passed through (see, for example, Patent Documents 1 to 4). However, even if the CTA slurry is sufficiently heated, clogging occurs in the manufacturing apparatus, and various problems may occur during continuous operation. For this reason, various methods have conventionally been proposed to solve these problems.

  For example, when undissolved terephthalic acid crystals flow into the packed bed of the hydrogenation catalyst, stable continuous operation may be difficult (see Patent Documents 1 and 4). In these patent documents, as a countermeasure, a retention zone formed by partitioning an overflow wall is provided at the upper inlet portion of the packed column reactor, and an aqueous CTA solution is supplied to the retention zone to overflow the overflow wall. After that, a method for producing PTA has been proposed in which liquid is passed through a catalyst packed bed located below the residence zone.

On the other hand, clogging may occur inside the heat exchanger, and stable continuous operation may be difficult. In particular, when heating is performed in stages using a plurality of heat exchangers, this problem has occurred, and an improvement measure has been demanded.
JP-A-6-128191 JP-A-8-208561 Japanese Patent Laid-Open No. 8-225489 JP 2004-203864 A

  The present invention is intended to solve the problems associated with the prior art as described above, and prevents the generation of clogging due to the scale formation and growth of terephthalic acid in the heat exchanger, thereby preventing a high concentration of crude terephthalate. An object of the present invention is to provide a method for producing high-purity terephthalic acid, which can form an acid aqueous solution and efficiently perform a hydrogenation reaction, thereby efficiently producing PTA.

  As a result of intensive studies to solve the above problems, the present inventors have compared the slurry temperature of 200 ° C. or less when clogging in the heat exchanger is heated stepwise using a plurality of heat exchangers. We found that this clogging was caused by terephthalic acid scale formation and growth. Furthermore, by controlling the flow rate of the CTA slurry having a temperature of 200 ° C. or lower in the heat exchanger to 1.7 m / sec or more, scale generation of terephthalic acid in the heat exchanger and clogging due to its growth do not occur. As a result, the present invention has been completed.

That is, the method for producing high-purity terephthalic acid according to the present invention comprises (I) a step of mixing crude terephthalic acid obtained by liquid phase oxidation of paraxylene and water to form a crude terephthalic acid slurry, (II) A step of heating and dissolving the crude terephthalic acid slurry stepwise in a plurality of heat exchangers to form a crude terephthalic acid aqueous solution, (III) a step of hydrogenating the crude terephthalic acid aqueous solution, and (IV) a step after hydrogenation In a method for producing high-purity terephthalic acid, including a step of stepwise cooling and cooling a terephthalic acid aqueous solution in a plurality of crystallization tanks to crystallize terephthalic acid, and (V) solid-liquid separation of the resulting terephthalic acid slurry. In the step (II), the flow rate of the crude terephthalic acid slurry having a temperature of 200 ° C. or less in the heat exchanger is 1.7 m / sec or more.

  The step (II) preferably includes at least one step (II-A) of heating the crude terephthalic acid slurry by introducing the steam generated by the pressure reduction cooling in the step (IV) into a heat exchanger. The high-temperature fluid containing the steam generated by the step-down cooling in the step (IV) and the condensate generated in the heat exchanger into which the steam at a higher temperature and pressure is introduced is exchanged at a lower temperature than the heat exchanger. It is more preferable to include at least one step (II-B) of introducing into a vessel and heating the crude terephthalic acid slurry.

  The heat exchanger is preferably a multi-tube heat exchanger, and the concentration of the crude terephthalic acid in the crude terephthalic acid slurry is preferably 10 to 40% by weight.

  According to the present invention, when a crude terephthalic acid slurry is heated to prepare a crude terephthalic acid aqueous solution, the generation of scales of terephthalic acid in a heat exchanger and the occurrence of clogging due to the growth thereof are prevented, and high concentration crude terephthalic acid An aqueous solution is formed and the hydrogenation reaction is performed efficiently, whereby PTA can be produced efficiently.

The method for producing high-purity terephthalic acid according to the present invention comprises (I) a step of mixing crude terephthalic acid obtained by liquid phase oxidation of paraxylene and water to form a crude terephthalic acid slurry, (II) the crude terephthalic acid A step of heating and dissolving the acid slurry stepwise in a plurality of heat exchangers to form a crude terephthalic acid aqueous solution, (III) a step of hydrogenating the crude terephthalic acid aqueous solution,
(IV) includes a step of stepwise cooling the crude terephthalic acid aqueous solution after hydrogenation in a plurality of crystallization tanks to crystallize terephthalic acid, and (V) a step of solid-liquid separation of the obtained terephthalic acid slurry. .

(I) Slurry forming step:
In this step, a crude terephthalic acid slurry is formed by mixing crude terephthalic acid obtained by liquid phase oxidation of para-xylene and water. The liquid phase oxidation of paraxylene can be carried out in an oxidation reactor generally used for the production of terephthalic acid. The oxidation reactor preferably contains a raw material such as para-xylene, a catalyst, and a solvent, and can perform continuous liquid phase oxidation by blowing air while replenishing these raw materials.

  The liquid phase oxidation of paraxylene is usually performed using a solvent and a catalyst. Examples of the solvent include fatty acids such as acetic acid, propionic acid, n-butyric acid, isobutyric acid, n-valeric acid, trimethylacetic acid, caproic acid, and mixtures of these with water. Of these, acetic acid or a mixed solvent of acetic acid and water is preferable.

  Examples of the catalyst include heavy metals, bromine, and compounds thereof. Examples of heavy metals include nickel, cobalt, iron, chromium, manganese, and the like. These catalysts are preferably used in the form of being dissolved in the reaction system. As said catalyst, it is preferable to use a cobalt compound, a manganese compound, and a bromine compound together, and the usage-amount of a cobalt compound is 10-10,000 ppm normally in conversion of cobalt with respect to a solvent, Preferably it is 100-3000 ppm. Moreover, the manganese compound is preferably 0.001 to 10 in terms of atomic ratio of manganese to cobalt, more preferably 0.001 to 2, and the bromine compound is preferably in terms of atomic ratio of bromine to cobalt, preferably 0.001 to 10. More preferably, it is 0.001-5.

  The liquid phase oxidation of para-xylene is usually performed using a molecular oxygen-containing gas. As such an oxygen-containing gas, a diluted oxygen gas obtained by diluting oxygen with an inert gas is usually used. For example, air or air enriched with oxygen is used. The temperature of the oxidation reaction is usually 150 to 270 ° C., preferably 170 to 220 ° C., and the pressure is at least the pressure at which the mixture can maintain a liquid phase at the reaction temperature, and usually 0.5 to 4 MPa (gauge pressure). . Further, although the reaction time depends on the size of the apparatus, etc., the residence time is usually about 20 minutes to 180 minutes. The water concentration in the reaction system is usually 3 to 30% by weight, preferably 5 to 15% by weight.

  CTA obtained by the liquid phase oxidation reaction is separated and recovered from the mother liquor in the liquid phase oxidation reaction by solid-liquid separation. The recovered CTA usually contains about 0.1 to 0.4% by weight of 4-CBA. The CTA and water are mixed in a mixing tank to form a CTA slurry. The said mixing tank can use what is generally used for manufacture of a terephthalic acid. The CTA concentration of the CTA slurry is usually 10 to 40% by weight, preferably 20 to 30% by weight.

(II) Aqueous solution formation process:
Terephthalic acid has low solubility in water at room temperature and normal pressure, and high temperature and pressure are required to increase the solubility in water. In this step, first, the CTA slurry is pressurized by a pressure pump, and then heated stepwise by a plurality of heat exchangers to dissolve the CTA to form a CTA aqueous solution. As a heat exchanger, what is generally used for manufacture of a terephthalic acid can be used, Preferably a multitubular heat exchanger is used.

  A high temperature fluid, particularly steam, is introduced into these heat exchangers, and the high temperature fluid and / or steam exchanges heat with the CTA slurry, whereby the CTA slurry is heated and CTA is dissolved to form a CTA aqueous solution. A high-temperature fluid and / or steam from other paths may be introduced into the heat exchanger, but steam generated by step-down cooling is introduced in the crystallization tank in step (IV) described later (hereinafter referred to as this The step is referred to as “step (II-A)”. Moreover, you may use together the process [process (II-C)] using other heaters, such as a heater. This step (II-C) may be provided as a preheating step before the step (II-A) and the step (II-B) described later, or the step (II-A). And after the process (II-B) mentioned later, you may be provided as a post-heating process.

Furthermore, in the present invention, at least one of the heat exchangers is introduced with steam generated by step-down cooling in a crystallization tank in step (IV) described later [step (II-A)], and the remaining At least one of the heat exchangers includes steam generated by step-down cooling in a crystallization tank in step (IV) described later, and condensate generated in a heat exchanger having a temperature higher than that of the heat exchanger. It is preferable that a high-temperature fluid is introduced [step (II-B)]. In this case, the steam and the high-temperature fluid exchange heat with the CTA slurry to heat the CTA slurry, and the steam and the high-temperature fluid become high-pressure condensate, and are usually used as a heat source for a lower-temperature heat exchanger. In particular, the condensate obtained by condensing the steam introduced into the heat exchanger in the step (II-A) is reduced in pressure, and then introduced into the steam generated in the crystallization tank in the step (IV) to introduce a high-temperature fluid. And is introduced into a heat exchanger having a temperature lower than that of the heat exchanger in the step (II-A). The pressure of the reduced condensate is approximately the same as the pressure of the steam introducing the condensate.

It is preferable that only one step (II-A) is included in step (II) from the viewpoint of energy efficiency.
The heating temperature in the steps (II-A) and (II-A) is determined by the temperature of the steam generated in step (IV). Further, the pressure of the introduced steam and the high-temperature fluid is also determined by the pressure of the steam generated in the step (IV).

  In this step, in order to sufficiently dissolve CTA, the CTA aqueous solution is usually heated to 230 ° C. or higher, preferably 240 to 300 ° C. The pressure in the system at the time of heating is not particularly limited as long as it is equal to or higher than the pressure at which the aqueous solution can be substantially maintained in the liquid phase, but is usually 1 to 11 MPa (gauge pressure), preferably 3 to 9 MPa (gauge pressure).

The concentration of the CTA aqueous solution thus prepared is determined by the concentration of the CTA slurry, and is usually 10 to 40% by weight, preferably 20 to 30% by weight.
In the present invention, in this step (II), the flow rate of the CTA slurry having a temperature of 200 ° C. or lower in the heat exchanger is 1.7 m / second or more, preferably 2.0 m / second or more, more preferably 2.2 m / second. Control over seconds. The upper limit of the flow rate is preferably 10.0 m / second or less and more preferably 5.5 m / second or less in order to sufficiently dissolve terephthalic acid. By controlling the flow rate of the CTA slurry having a temperature of 200 ° C. or less in the heat exchanger within the above range, clogging due to scale generation in the heat exchanger can be prevented, and an aqueous CTA solution can be efficiently prepared.

(III) Hydrogenation process:
In this step, the CTA aqueous solution is introduced into a hydrogenation reactor and subjected to hydrogenation treatment, and 4-CBA contained in the CTA aqueous solution is reduced to 4-methylbenzoic acid (4-MBA).

  If the said hydrogenation reaction tank has a catalyst layer filled with the hydrogenation catalyst and can supply hydrogen in the state which the catalyst and CTA aqueous solution contacted, the shape, structure, etc. will not be restrict | limited in particular. A preferable hydrogenation reaction tank has a fixed catalyst layer filled with a solid catalyst inside, an aqueous solution introduction path and an aqueous solution outlet path so that a CTA aqueous solution can be passed through the catalyst layer, and hydrogen is further supplied. The thing which has a hydrogen supply path so that it can supply is mentioned. The flow direction of the CTA aqueous solution is not limited and may be an upward flow or a downward flow. However, it is preferable that the CTA aqueous solution flow in the downward flow. For this purpose, the aqueous solution introduction path is formed above the hydrogenation reaction tank. It is preferable that the path is provided in the lower part. Moreover, it is preferable that hydrogen is supplied from the upper part, and for this purpose, it is preferable that a hydrogen supply path is provided in the upper part of the hydrogenation reaction tank.

  As the hydrogenation catalyst, those conventionally used can be used, for example, palladium, ruthenium, rhodium, osmium, iridium, platinum, platinum black, palladium black, iron, cobalt-nickel, etc. A solid catalyst in which these are supported on an adsorbent carrier such as activated carbon so that they can be formed is preferable.

As a specific hydrogenation treatment method, for example, a CTA aqueous solution is introduced into a hydrogenation reaction tank, and while passing through the catalyst layer, hydrogen gas is usually 1.5 times mol or more of 4-CBA in the CTA aqueous solution, preferably It is desirable to perform hydrogenation by supplying at a flow rate of 2 times mole or more. The reaction temperature at the time of hydrogenation is usually 230 ° C. or more, preferably 255 to 300 ° C., the pressure is usually 1 to 11 MPa, preferably 3 to 9 MPa (gauge pressure), and the hydrogen partial pressure is usually 0.05 MPa or more. Preferably it is 0.05-2 MPa.
By this hydrogenation treatment, 4-CBA in CTA is reduced to water-soluble 4-MBA, while TA is hardly soluble in water, so that it is usually at a temperature of 300 ° C. or less, preferably 100 to 280 ° C. By performing precipitation and solid-liquid separation, 4-MBA can be separated from the CTA aqueous solution to obtain PTA.

(IV) Crystallization process:
In this step, TA is crystallized by stepwise cooling the CTA aqueous solution after hydrogenation (hereinafter referred to as “CTA hydrogenation solution”) stepwise in a plurality of crystallization tanks. As the crystallization tank, those generally used for the production of terephthalic acid can be used. Specifically, the CTA hydrogenation treatment liquid is introduced into a crystallization tank set to a pressure condition lower than the pressure of the CTA hydrogenation treatment liquid, and the pressure of the CTA hydrogenation treatment liquid is reduced. This is a method of cooling the treatment liquid (step-down cooling). In the present invention, the pressure of the CTA hydrogenation treatment liquid is reduced stepwise using a plurality of crystallization tanks, and the temperature of the CTA hydrogenation treatment liquid is lowered stepwise accordingly. Thereby, terephthalic acid is crystallized and TA slurry is formed.

  When the step-down cooling is performed as described above, a part of the aqueous medium in the CTA hydrogenation treatment liquid is vaporized and vapor is generated in each crystallization tank. As described above, this steam is preferably introduced into the heat exchanger in the step (II) and used as a heating source for the CTA slurry. In particular, among the above crystallization tanks, the steam generated by the pressure reduction cooling in the highest temperature and high pressure crystallization tank alone is not introduced into the heat exchanger of the step (II) without introducing the condensate generated in the heat exchanger. It is preferable to introduce from the viewpoint of energy efficiency.

The temperature of the CTA hydrogenation treatment liquid cooled by the step-down cooling and the generated steam are appropriately determined by adjusting the pressure during the step-down cooling in each crystallization tank.
(V) Solid-liquid separation process:
In this step, the TA slurry is solid-liquid separated to separate and recover PTA from the mother liquor. Solid-liquid separation can be carried out using a solid-liquid separation apparatus such as a filter or a centrifuge, which is generally used for producing terephthalic acid.

Furthermore, the PTA may be suspended again in water, and the foreign matter adhering to the PTA crystal may be transferred to the aqueous phase, followed by solid-liquid separation and drying.
(Manufacturing equipment)
The method for producing high-purity terephthalic acid according to the present invention comprises a mixing tank, a plurality of heat exchangers, a reaction tank, a plurality of crystallization tanks, and a solid-liquid separation device connected in series in this order. It is a manufacturing apparatus, Comprising: It can carry out using the following manufacturing apparatuses which have n heat exchangers and n crystallization tanks.

<First manufacturing apparatus>
In the first manufacturing apparatus, the heat exchanger and the crystallization tank are not connected, and steam generated in the crystallization tank is not used. In this case, steam generated in another route is introduced into the heat exchanger (see FIG. 1 (when n = 4)).

<Second manufacturing apparatus>
In the second manufacturing apparatus, at least one heat exchanger is connected to the crystallization tank so that steam generated in the crystallization tank is introduced. The remaining heat exchangers may be connected to other heat source introduction paths, or may be heated by other heaters such as a heater, but all heat exchangers are connected to the crystallization layer. Is preferable from the viewpoint of energy efficiency (see FIG. 2 (when n = 4)).

  When multiple heat exchangers are connected to the crystallization tank, the highest temperature heat exchanger among these heat exchangers is connected to the highest temperature and pressure crystallization tank. It is preferable from the viewpoint.

<Third and fourth manufacturing apparatuses>
In the third and fourth production apparatuses, at least one heat exchanger is connected to the crystallization tank so that steam generated in the crystallization tank is introduced.

  Further, at least one of the remaining heat exchangers is connected to the crystallization tank by a steam pipe, and is connected to a heat exchanger having a higher temperature and a higher pressure than the heat exchanger in the middle of the steam pipe. A condensate pipe is connected, and a pressure regulating valve is provided in the middle of the condensate pipe. Thereby, the condensate produced | generated with a heat exchanger is introduce | transduced into the vapor | steam which generate | occur | produces in a crystallization tank, a high temperature fluid is formed, and this high temperature fluid is introduce | transduced into a heat exchanger of low temperature and low pressure rather than this heat exchanger. Such 3rd and 4th manufacturing apparatus is preferable at the point which is excellent in energy efficiency compared with the said 1st and 2nd manufacturing apparatus.

Of the heat exchangers, the highest temperature heat exchanger is preferably connected to the highest temperature and pressure crystallization tank from the viewpoint of energy efficiency.
In addition, at least one of the heat exchangers (hereinafter, this heat exchanger is referred to as “heat exchanger α”) is connected to a crystallization tank (hereinafter, this crystallization tank is referred to as “crystallization tank α”), At least one of the heat exchangers at a lower temperature and lower pressure than the heat exchanger α (hereinafter, this heat exchanger is referred to as “heat exchanger β”) has a crystallization tank (hereinafter referred to as this crystallization tank) different from the crystallization tank α. The crystallization tank is connected to the crystallization tank β) by a steam pipe, and a condensate pipe connected to the heat exchanger α is connected in the middle of the steam pipe. It is preferable that a pressure regulating valve is provided.

Furthermore, it is preferable from the viewpoint of energy efficiency that the production apparatus of the present invention has one heat exchanger connected only to the crystallization tank.
Specific examples of such a manufacturing apparatus (when n = 4) are shown in FIGS. In the third manufacturing apparatus shown in FIG. 3, the heat exchanger (n) is connected to the crystallization tank (1) by a pipe, and the i-th heat exchange (i is an integer of 1 to n-1) from the low temperature side. The apparatus (i) is connected to the crystallization tank (n−i + 1) by a steam pipe, and a condensate pipe connected to the heat exchanger (i + 1) is connected to the steam pipe (see FIG. 3). ). The condensate pipe is preferably provided with a pressure adjusting valve.

  In the fourth manufacturing apparatus shown in FIG. 4, the heat exchanger (n) is connected to the crystallization tank (1) by piping, and the heat exchanger (i) is connected to the crystallization tank (n−i + 1) by steam piping. In addition, a condensate pipe connected to the heat exchanger (i + 2) is connected in the middle of the steam pipe (where i is an integer of 1 to n-1). Also in this case, it is preferable that the condensate pipe is provided with a pressure adjusting valve.

  Hereinafter, in the case of producing high-purity terephthalic acid with the third or fourth production apparatus, the vapor generated in the crystallization tank and the method of using the condensate thereof, specifically, in step (IV), the crystallization tank is Explains how to use n heat exchangers (n is an integer of 2 or more) and n heat exchangers that use steam generated from the n crystallization tanks as a heat source in step (II). To do.

The highest temperature and pressure crystallization tank is the crystallization tank (1), and the crystallization tank (2), the crystallization tank (3),. Let the low-temperature and low-pressure crystallization tank be a crystallization tank (n). Moreover, let the heat exchanger (1) the lowest temperature be a heat exchanger (1), and in order of the high temperature side, heat exchanger (2), heat exchanger (3), ..., heat exchanger (n-1), The hottest heat exchanger is defined as a heat exchanger (n).

  Normally, the steam generated by the pressure reduction cooling in the highest temperature and high pressure crystallization tank (1) is introduced into the heat exchanger (n) alone without introducing the condensate produced in the heat exchanger from the viewpoint of energy efficiency. It is preferable to introduce.

  In addition, the i-th heat exchanger (i) from the low temperature side (i is an integer of 1 to n-1) has a high pressure generated by the heat exchanger (i + 1) that is higher in temperature than the heat exchanger (i). A high-temperature fluid formed by introducing the condensate into steam generated by pressure reduction cooling in the i-th crystallization tank from the low temperature side is introduced as a heating source. Here, the i-th crystallization tank from the low temperature side is the (n−i + 1) th when counted from the high temperature side. Therefore, the i-th heat exchanger (i) from the low temperature side is introduced with a high-temperature fluid of the steam generated in the crystallization tank (n−i + 1) and the high-pressure condensate generated in the heat exchanger (i + 1). The When forming this high-temperature fluid, it is preferable to introduce the condensate into a steam pipe connected to a low-temperature and low-temperature heat exchanger after reducing the pressure of the condensate than the heat exchanger in which the condensate was generated. .

The former process corresponds to the process (II-A), and the latter process corresponds to the process (II-B).
In the production method of the present invention, steps (II-A) and (II-B) are not limited to the above steps, and for example, step (II-B) may be the following steps.

  The i-th heat exchanger (i) from the low temperature side (i is an integer of 1 to n-1) has a high-pressure condensate generated in the heat exchanger (i + 2) that is higher in temperature than the heat exchanger (i). May be introduced as a heating source by introducing a high-temperature fluid into the steam generated by the pressure reduction cooling in the i-th crystallization tank from the low temperature side. Similarly to the above, since the i-th crystallization tank from the low temperature side is the (ni + 1) th when counted from the high temperature side, the i-th heat exchanger (i) from the low temperature side has the crystallization tank (n-i + 1). ) And a high-temperature fluid of high-pressure condensate generated in the heat exchanger (i + 2) may be introduced. In this case as well, when forming the high-temperature fluid, after reducing the pressure of the condensate, the condensate is put into a steam pipe connected to a heat exchanger having a lower pressure and lower temperature than the heat exchanger in which the condensate is generated. It is preferable to introduce.

[Example]
EXAMPLES Hereinafter, although an Example demonstrates this invention, this invention is not limited at all by this Example.

  Using a production apparatus (n = 4) having four multi-tube heat exchangers and four crystallization tanks shown in FIG. 2, a CTA slurry having a CTA concentration of 30% by weight is shown in Table 1. Then, the resulting CTA aqueous solution was subjected to hydrogenation treatment, and the CTA treatment solution was cooled under reduced pressure to produce high-purity terephthalic acid. A flow meter and a flow control valve are installed in front of the first heat exchanger (i = 1), and the flow rate of the CTA slurry is measured with this flow meter. The flow rate of the CTA slurry is determined from the flow rate and the capacity of the heat exchanger. The flow rate was calculated. In addition, high-purity terephthalic acid was produced for 2 months each under different flow rate conditions.

Steam having a temperature of 143 ° C. and a pressure of 0.4 MPa generated in the fourth crystallization tank (i = 4) was introduced into the first heat exchanger (i = 1). Steam having a temperature of 182 ° C. and a pressure of 1.1 MPa generated in the third crystallization tank (i = 3) was introduced into the second heat exchanger (i = 2). Steam having a temperature of 207 ° C. and a pressure of 1.8 MPa generated in the second crystallization tank (i = 2) was introduced into the third heat exchanger (i = 3). Steam at a temperature of 220 ° C. and a pressure of 2.3 MPa generated in the first crystallization tank (i = 1) was introduced into the fourth heat exchanger (i = 4). The condensate after heat exchange in each heat exchanger has a temperature of 1 respectively.
The pressure was reduced to 00 ° C. and the pressure to 0.1 MPa, and then recovered in a condensate recovery tank.

  At any flow rate, the temperature of the CTA slurry at the outlet of the first heat exchanger is 130 ° C to 135 ° C, the temperature of the CTA slurry at the outlet of the second heat exchanger is 180 to 185 ° C, and the third heat The temperature of the CTA slurry at the exchanger outlet was 200 to 205 ° C.

  Table 1 shows the pressure loss increase rate of the CTA slurry from the first heat exchanger to the second heat exchanger. A slight increase in pressure loss was observed, but no clogging in the heat exchanger was observed.

  Moreover, the energy efficiency in the said manufacturing method was evaluated by the exergy loss produced when depressurizing the condensate produced | generated with the heat exchanger. The results are shown in Table 2.

  Using a production apparatus (n = 4) having four multi-tube heat exchangers and four crystallization tanks shown in FIG. 3, a CTA slurry having a CTA concentration of 30% by weight is shown in Table 3. Then, the resulting CTA aqueous solution was subjected to hydrogenation treatment, and the CTA treatment solution was cooled under reduced pressure to produce high-purity terephthalic acid for one month. The flow rate of the CTA slurry was calculated in the same manner as in Example 1.

In the first heat exchanger (i = 1), the temperature of the steam generated in the fourth crystallization tank (i = 4) is the condensate after heat exchange in the second heat exchanger (i = 2) (143 C.) and a pressure (0.4 MPa), the pressure was reduced to be equal to the pressure (0.4 MPa), and then introduced into the steam generated in the fourth crystallization tank to form a high-temperature fluid, and this high-temperature fluid was introduced. In the second heat exchanger (i = 2), the steam (182 ° C.) generated in the third crystallization tank (i = 3) is the condensate after heat exchange in the third heat exchanger (i = 3). After the pressure was reduced to be equal to the temperature and pressure (1.1 MPa), a high-temperature fluid was formed by introducing into the steam generated in the third crystallization tank, and this high-temperature fluid was introduced. In the third heat exchanger (i = 3), the condensate after the heat exchange in the fourth heat exchanger (i = 4) is the temperature of the steam generated in the second crystallization tank (i = 2) (207 C.) and a pressure (1.8 MPa), the pressure was reduced to be equal to the pressure, and the high temperature fluid was formed by introducing the vapor generated in the second crystallization tank to introduce the high temperature fluid. Steam having a temperature of 220 ° C. and a pressure of 2.3 MPa, which was generated in the first crystallization tank (i = 1), was introduced into the fourth heat exchanger (i = 4). The condensate produced in the first heat exchanger was pressure-reduced so that the temperature was 100 ° C. and the pressure was 0.1 MPa, and then recovered in the condensate recovery tank.

  At any flow rate, the temperature of the CTA slurry at the outlet of the first heat exchanger is 120 to 125 ° C., and the temperature of the CTA slurry at the outlet of the second heat exchanger is 160 to 165 ° C. The temperature of the CTA slurry at the outlet of the vessel was 200 to 205 ° C.

  Table 3 shows the pressure loss increase rate of the CTA slurry from the first heat exchanger to the second heat exchanger. A slight increase in pressure loss was observed, but no clogging in the heat exchanger was observed.

  Moreover, the energy efficiency in the said manufacturing method was evaluated by the exergy loss produced when depressurizing the condensate produced | generated with the heat exchanger. The results are shown in Table 4.

  Using a production apparatus (n = 4) having four multi-tubular heat exchangers and four crystallization tanks shown in FIG. 3, a CTA slurry having a CTA concentration of 30% by weight is obtained at a flow rate shown in Table 5. The mixture was supplied to a heat exchanger and heated as described below. The resulting CTA aqueous solution was subjected to hydrogenation treatment, and then the CTA treatment solution was cooled down to produce high purity terephthalic acid. The flow rate of the CTA slurry was calculated in the same manner as in Example 1. High purity terephthalic acid was produced for 2 months each under different flow rate conditions.

  In the first heat exchanger (i = 1), the temperature of the steam generated in the fourth crystallization tank (i = 4) is the condensate after heat exchange in the second heat exchanger (i = 2) (143 C.) and a pressure (0.4 MPa), the pressure was reduced to be equal to the pressure (0.4 MPa), and then introduced into the steam generated in the fourth crystallization tank to form a high-temperature fluid, and this high-temperature fluid was introduced. In the second heat exchanger (i = 2), the steam (182 ° C.) generated in the third crystallization tank (i = 3) is the condensate after heat exchange in the third heat exchanger (i = 3). After the pressure was reduced to be equal to the temperature and pressure (1.1 MPa), a high-temperature fluid was formed by introducing into the steam generated in the third crystallization tank, and this high-temperature fluid was introduced. In the third heat exchanger (i = 3), the condensate after the heat exchange in the fourth heat exchanger (i = 4) is the temperature of the steam generated in the second crystallization tank (i = 2) (207 C.) and a pressure (1.8 MPa), the pressure was reduced to be equal to the pressure, and the high temperature fluid was formed by introducing the vapor generated in the second crystallization tank to introduce the high temperature fluid. Steam having a temperature of 220 ° C. and a pressure of 2.3 MPa, which was generated in the first crystallization tank (i = 1), was introduced into the fourth heat exchanger (i = 4). The condensate generated in the first heat exchanger was reduced in pressure so that the temperature was 100 ° C. and the pressure was 0.1 MPa, and then recovered in the condensate recovery tank.

  At any flow rate, the temperature of the CTA slurry at the outlet of the first heat exchanger is 115 to 120 ° C., the temperature of the CTA slurry at the outlet of the second heat exchanger is 130 to 135 ° C., and the third heat exchange is performed. The temperature of the CTA slurry at the outlet of the vessel was 200 to 205 ° C.

  Table 5 shows the pressure loss increase rate of the CTA slurry from the first heat exchanger to the second heat exchanger. A slight increase in pressure loss was observed, but no clogging in the heat exchanger was observed.

Moreover, when the energy efficiency in the said manufacturing method was evaluated by the exergy loss produced when depressurizing the condensate produced | generated with the heat exchanger, it was the same as Example 2.
[Comparative Example 1]
Except that the flow rate of the CTA slurry in the heat exchanger was changed to 1.6 m / sec, high purity terephthalic acid was produced in the same manner as in Example 1, and the CTA slurry was added to the heat exchanger in less than one month. It became difficult to introduce. When the inside of the heat exchanger was inspected, clogging occurred inside.

FIG. 1 is a schematic view of a first manufacturing apparatus (when n = 4). FIG. 2 is a schematic view of the second manufacturing apparatus (when n = 4). FIG. 3 is a schematic view of a third manufacturing apparatus (when n = 4). FIG. 4 is a schematic view of a fourth manufacturing apparatus (when n = 4).

Explanation of symbols

1 Mixing tank 21 1st multi-tube heat exchanger (i = 1)
22 Second multi-tube heat exchanger (i = 2)
23 Third multi-tube heat exchanger (i = 3)
24 Fourth multi-tube heat exchanger (i = 4)
3 Heat exchanger 4 Hydrogenation reaction tank 51 First crystallization tank (i = 1)
52 Second crystallization tank (i = 2)
53 Third crystallization tank (i = 3)
54 Fourth crystallization tank (i = 4)
6 Solid-liquid separators 71 to 74 Pressure regulating valve 8 Condensate recovery tank 9 Steam pipe 10 Condensate pipe CTA Crude terephthalic acid PTA High purity terephthalic acid

Claims (5)

(I) A step of mixing crude terephthalic acid obtained by liquid phase oxidation of para-xylene and water to form a crude terephthalic acid slurry, and (II) heating the crude terephthalic acid slurry in stages with a plurality of heat exchangers. (III) Step of hydrogenating the crude terephthalic acid aqueous solution, (IV) Step-by-step cooling of the aqueous terephthalic acid solution after hydrogenation in multiple crystallization tanks In the method for producing high-purity terephthalic acid, including the step of crystallizing terephthalic acid, and (V) solid-liquid separation of the obtained terephthalic acid slurry,
A process for producing high-purity terephthalic acid, characterized in that, in the step (II), the flow rate of the crude terephthalic acid slurry having a temperature of 200 ° C. or less in the heat exchanger is 1.7 m / sec or more.   The step (II) includes at least one step (II-A) of heating the crude terephthalic acid slurry by introducing the steam generated by the pressure reduction cooling in the step (IV) into a heat exchanger. The manufacturing method of the high purity terephthalic acid of Claim 1.   In the step (II), a high-temperature fluid containing the steam generated by the step-down cooling in the step (IV) and a condensate generated in a heat exchanger into which steam having a temperature higher than that of the steam is introduced is added to the heat. A method for producing high-purity terephthalic acid, comprising at least one step (II-B) of introducing the crude terephthalic acid slurry by introducing it into a heat exchanger having a temperature lower than that of the exchanger.   The method for producing high-purity terephthalic acid according to any one of claims 1 to 3, wherein the heat exchanger is a multitubular heat exchanger.   The method for producing high-purity terephthalic acid according to any one of claims 1 to 4, wherein a concentration of the crude terephthalic acid in the crude terephthalic acid slurry is 10 to 40% by weight.

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