CN1634848A - Process for producing terephthalic acid and device therefor - Google Patents

Abstract

The invention discloses a production method and device for producing terephthalic acid (TA). The device consists of three parts: bubble column reactor, slurry treatment device, and tail gas treatment device; the overall idea of the method is: use a bubble column reactor with a gas separation section to realize liquid-phase catalytic oxidation of p-xylene; The slurry treatment device adopts multi-stage evaporation, filtration and drying steps to recover solid terephthalic acid; the tail gas treatment device adopts multi-stage condensation by-product steam to recover reaction heat, and adopts azeotropic distillation method to recover solvent; each unit, equipment, The configuration of the flow stream and related process conditions are optimized and determined according to the minimum production cost of the product. The invention can optimize flow design and process improvement, so that the manufacturing cost is lower and the energy saving effect is better.

Description

Method and device for producing terephthalic acid

technical field

The invention relates to a production method and device for producing terephthalic acid (TA).

Background technique

Terephthalic acid is an important raw material for the production of polyester (PET) fibers and resins. At present, it is mainly produced by the p-xylene air oxidation method. Or tetrabromoethane) in the acetic acid solvent, pass into air or enriched oxygen to carry out oxidation, generate solid product terephthalic acid. The typical reaction temperature is 155-205°C, the pressure is 0.5-1.6MPa, and the residence time is 40-120min. The reaction heat is removed by solvent evaporation, and the steam is condensed and returned to the reactor. The resulting slurry is then obtained through subsequent separation and refining processes. Purified terephthalic acid (PTA) products.

The production process of TA includes three main parts: reaction unit, tail gas treatment unit and slurry treatment unit. The tasks and functions of each unit can be represented by the block diagram in Figure 1. The raw materials p-xylene and air are fed into the bubble column reactor, and the liquid-phase oxidation occurs under the action of the catalyst to generate solid terephthalic acid. After the slurry is discharged from the bottom of the reactor, it enters the slurry treatment unit. The main task of this unit is to separate the mother liquor and TA solids, recover the energy in the slurry, purify the mother liquor to control the concentration of impurities, and recover the catalyst at the same time. The heat of reaction is taken out by the evaporation of the solvent acetic acid, and the acetic acid vapor and tail gas move out of the reactor from the top and enter the tail gas treatment unit. The main task of this unit is to recover reaction heat, recover acetic acid solvent, and purify tail gas.

Due to the large amount of heat released during the oxidation of p-xylene (1.26×10 ⁷ J of reaction heat is released after oxidation of p-xylene per kilogram), and the energy consumption of each unit and equipment in the process is also high, so the comprehensive utilization of energy is a very important problems that require careful consideration in reactor design and process design. First of all, in terms of reactor design, the current industrialized bubble column reactor mainly adopts a tank structure with stirring paddles, such as the stirred tank reactors introduced in patent US5211924 and patent US5102630 with two layers of stirring paddles. This type of reactor equipment is expensive and consumes a lot of electricity during operation (2-3kw per cubic volume), and the heat of reaction has not been directly and effectively utilized. In the reactors disclosed in patent ZL94103145.4 and US5463113, a rectification tower is arranged above the oxidizer, and the heat of reaction is used as the heat source of rectification to make energy utilization more reasonable. However, the upper rectifying section of this reactor is too slender, so that the height of the lower reaction section is limited, and the ratio of height to diameter of the reaction section is small. During industrial implementation, it is often necessary to install a stirring paddle upwards at the bottom to promote mixing and strengthen solid suspension. It increases the reactor investment and operating costs.

Secondly, in terms of process design, the existing process generally adopts the method of multi-stage condensation of tail gas to generate steam of different energy levels to recover the heat of reaction, and a part of the condensate is extracted for rectification and dehydration to recover acetic acid solvent; for the treatment of reaction slurry , Some processes use multi-stage evaporation and filtration methods for liquid-solid separation, and some use multi-stage centrifugation methods for separation. The energy consumption indicators of different processes are different, but there is still great potential in terms of energy saving. In order to further save energy and tap the potential, the patent US5723656 discloses a method that directly introduces the tail gas at the top of the bubble column reactor into a rectification tower, and then performs multi-stage condensation and heat exchange on the tail gas at the top of the rectification tower to recover its energy, and the tail gas then enters the expansion The machine does work. The method introduced in the patent US6143925 is to install a high-efficiency rectification tower above the reactor to separate acetic acid and water, and the tail gas is condensed in multiple stages to generate steam of different energy levels, and then the steam is introduced into the turbine to perform work. The method introduced in patent ZL99108695.3 and ZL99109477.8 is to connect the motor, compressor, steam turbine and expander coaxially, and the steam turbine is driven by the by-product steam of different energy levels, and the reaction tail gas is heated and then introduced into the expander to do work. But above-mentioned method is somewhat difficult to implement, and for example patent US5723656 and US6143925 introduce and finish the task of all acetic acid dehydration by a rectification tower that is connected with bubble column reactor and just require separation equipment very huge, because the temperature pressure of reaction tail gas is higher , but also contains a large amount of inert gas. On the other hand, the various patented technologies currently disclosed are partial energy-saving measures. Although they may produce certain energy-saving effects on some units and equipment after implementation, it is difficult to achieve the goal of the lowest energy consumption in the entire process. Therefore, it is necessary to use the method of system engineering to comprehensively consider and optimize the design of the energy integration of the whole system to form the best process design scheme.

Contents of the invention

The purpose of the present invention is to provide a terephthalic acid production method and equipment with lower cost and better energy-saving effect, which are used for process flow design and process improvement. In order to achieve the above purpose, the device consists of three parts: a bubble column reactor, a slurry treatment device, and a tail gas treatment device; Liquid-phase catalytic oxidation; the slurry treatment device adopts multi-stage evaporation, filtration and drying steps to recover solid terephthalic acid; the tail gas treatment device adopts multi-stage condensation by-product steam to recover reaction heat, and adopts azeotropic distillation method to recover solvent; in the process The configuration of each unit, equipment, stream and related process conditions are optimized and determined according to the minimum production cost of the product.

The invention provides a terephthalic acid production device, which includes a bubble column reactor, a slurry treatment device, and a tail gas treatment device. The bubble column reactor is in the shape of a straight cylinder, and the interior is a gas treatment section from top to bottom. The space section and the three-phase reaction section, the gas treatment section accounts for 10-30% of the total height of the bubble column reactor, and the three-phase reaction section accounts for 60-85% of the total height of the bubble column reactor; Packing or liquid holding tray, a gas distributor is installed horizontally at the lower part of the three-phase reaction section, the gas distributor is connected to the gas inlet of the bubble column, and a raw material feeding pipe is arranged on the three-phase reaction section above the gas distributor , There is a discharge pipe at the bottom of the tower, and a tail gas pipe and a condensate return pipe at the top of the tower; the air inlet of the bubble tower is connected with an air compressor pump. The slurry treatment device includes a multi-stage evaporator connected with the discharge pipe. After the upper outlet of the multi-stage evaporator passes through a heat exchanger, it is respectively connected with the acetic acid storage tank, the tail gas absorption tower, and the rectification tower. The lower outlet is connected with the filter, and the outlet of the filter is respectively connected with the raw material feeding pipe, the mother liquor purification device and the dryer; the outlet of the mother liquor purification device is connected with the acetic acid storage tank. The tail gas treatment device includes a multi-stage condenser. The shell side of the multi-stage condenser is connected to the steam turbine and heat source. The tube side of the multi-stage condenser passes through the vapor-liquid separator and is respectively connected to the condensate return pipe, rectification tower, and tail gas absorption. The towers are connected; the rectification tower is connected with the outlet of the heat exchanger, the lower outlet of the rectification tower is connected with the raw material feed pipe and the acetic acid storage tank, and the upper outlet of the rectification tower is connected with the condensate return pipe and the oil-water separator respectively. The outlet of the oil-water separator is respectively connected with the entrainer recovery tower and the p-xylene recovery tower; the outlet of the tail gas absorption tower is connected with the tail gas purifier.

Adopt the method for above-mentioned device to produce terephthalic acid, comprise the steps:

1) Oxidation reaction: p-xylene and solvent acetic acid enter the three-phase reaction section in the bubble column reactor respectively from the raw material feed pipe, and the oxygen-containing gas passes through the gas distributor from the gas inlet of the bubble column reactor for oxidation reaction ; The gas obtained from the reaction is discharged from the tail gas pipe after passing through the gas treatment section, and the slurry obtained from the reaction is discharged from the discharge pipe;

2) Slurry treatment: the slurry discharged from the discharge pipe is decompressed and cooled in the multi-stage evaporator, the gas discharged from the upper outlet of the multi-stage evaporator enters the heat exchanger for heat exchange, and the obtained condensate is introduced into acetic acid The storage tank, the non-condensable gas is introduced into the tail gas absorption tower for treatment, the steam is introduced into the rectification tower for dehydration and provides part of the heat source; the slurry discharged from the lower outlet of the multi-stage evaporator is filtered into mother liquor and solid by the filter, and the solid enters Dried in the dryer to recover acetic acid, most of the mother liquor returns to the three-phase reaction section through the raw material feeding pipe, a small part of the mother liquor is purified through the mother liquor purification device, and the metal catalyst components in the mother liquor are recovered at the same time, and the purified mother liquor enters Acetic acid storage tank;

3) Tail gas treatment: After the tail gas discharged from the tail gas pipe is processed by a multi-stage condenser, the generated steam drives the steam turbine through the shell side and provides heat source for the rectification tower, and the generated cooling water condenses the tail gas through the shell side; The condensed tail gas comes out from the tube side of the multi-stage condenser, and after passing through the vapor-liquid separator, the tail gas is purified successively through the tail gas absorption tower and the tail gas purifier, and part of the condensate is returned to the inside of the bubble column reactor through the condensate return pipe, and part of the condensate is It is introduced into the rectification tower to separate acetic acid and water, and the heat exchanger is also introduced into the rectification tower to separate acetic acid and water; part of the separated acetic acid enters the interior of the bubble column reactor through the raw material feeding pipe, and part of it enters the acetic acid storage tank; The condensate obtained by condensing the vapor from the outlet of the distillation tower enters the bubble column reactor through the reflux pipe partly, and partly enters the oil-water separator for separation. The separated oil phase is then introduced into the entrainer recovery tower for recovery of the azeotrope The column carries out rectification and dehydration, and the entrainer enters the p-xylene recovery tower to purify the entrainer.

Main points of the present invention are described in detail below:

1. Oxidation reaction unit

The structure of the bubble column reactor provided by the present invention is shown in Fig. 2, and the reactor is a straight cylindrical bubble column reactor with a gas separation section, and the inside is composed of three sections from top to bottom: the top is a gas treatment section , the middle and lower part is the three-phase reaction section, and the free space section is between the two. When working, the raw material p-xylene, solvent acetic acid and catalyst are mixed and put into the reactor from the raw material feeding pipe, and the oxygen-containing gas is bubbled in from the air inlet of the bubble tower through the gas distributor to react with p-xylene. This produces solid terephthalic acid. The solid slurry produced by the reaction is discharged from the discharge pipe and enters the slurry processing unit. After the mother liquor is filtered and separated, most of it flows back into the reactor, and a small part is drawn out for purification and catalyst recovery. The heat of reaction is removed by the evaporation of the solvent acetic acid. The solvent vapor and tail gas flow out of the reactor through the tail gas pipe and enter the tail gas recovery unit. Water, control the water content in the reactor.

The design points of bubble column reactor structure and process conditions include the following aspects:

1. Three-phase reaction section of bubble column reactor

The three-phase reaction section is the main part of the bubble column reactor where gas-liquid contact and chemical reactions take place. The bubble column reactor used in the present invention is a bubble column configuration, which is different from the stirred tank configuration used in most existing processes. The reason for choosing the bubble column is that according to chemical engineering research, the maximum reaction rate of p-xylene liquid phase oxidation is slower than the maximum gas-liquid mass transfer rate, for example, for a stirred tank, the ratio of the two rates is 0.2 to 0.3, For the bubble column, this ratio is between 0.3 and 0.4. Therefore, the industrial oxidation process is basically controlled by the liquid phase chemical reaction, and the configuration of the bubble column can meet the requirements of the reaction for gas-liquid mass transfer. At the same time, under the reaction conditions, the turbulence in the bubble column is violent and the mixing is good. It can also meet the requirements of the oxidation reaction process in terms of liquid phase mixing, solid suspension, solvent evaporation, and liquid phase oxygen concentration distribution. The obtained product is similar to that of a stirred tank. response effect. In addition, the structure of the bubble tower is simple, the cost is low, there are no moving parts such as stirring paddles, and the operation is more energy-saving and power-saving.

The height of the three-phase reaction section is the liquid level of the reaction slurry, accounting for 60-85% of the total height of the bubble column reactor, so as to ensure that the reaction section has sufficient volume and slurry residence time. The gas distributor is installed in the lower part of the tower. P-xylene is mixed with the solvent acetic acid, catalyst and refluxed mother liquor, and then added through the feed pipe from above the gas distributor at 1/4 to 1/2 of the height of the reaction section, and the slurry generated by the reaction is discharged from the slurry pipe at the bottom of the tower .

The gas distributor is a device for uniform distribution of gas, and its form can be varied. For example, conventional perforated plate distributors, annular distributors or multi-tube distributors can be used.

2. Gas separation section

There is a gas separation section on the upper part of the bubble column reactor, its function is mainly to make the rising steam and tail gas in the lower reaction section fully contact with the condensed and refluxed aqueous acetic acid at the top of the tower, on the one hand recover the solid particles entrained in the tail gas, on the other hand Carry out solvent distillation to separate acetic acid and water. The structure of the gas separation section is similar to that of a general rectification tower. It can use a high-efficiency packed tower, a sieve tray tower or a valve tower, or a combination of several trays and packing. The refluxed condensate flows down from the top of the tower, and the reaction tail gas and steam flow up, and the two contact in countercurrent to realize the recovery of solids and the separation of solvents.

Conventional bubble column reactors are usually equipped with an enlarged section in the upper part to separate solid particles and liquid foam, but this scheme is difficult to work for the p-xylene oxidation process, because the process has a large amount of gas treatment and solvent evaporation, and the empty column The gas velocity is extremely high (up to 0.3-1.0 m/s), and it is difficult to effectively separate solid particles and reduce liquid foam entrainment by using the expansion section. A gas-handling section consisting of trays or packing used to separate solids. On the other hand, setting a rectification section above the bubble column reactor can directly use the heat of reaction to separate the solvent acetic acid from water, which is beneficial to energy saving. Theoretically, the more the number of trays in the gas separation section, the higher the energy recovery rate. Therefore, patents US5723656 and US6143925 propose to use a high-efficiency rectification tower connected to the reactor to complete the task of recovering all acetic acid from the tail gas. But this is relatively difficult, because the temperature and pressure of the reaction are relatively high, and the tail gas also contains a large amount of non-condensable gas, so it is difficult to completely separate acetic acid and water. Therefore, considering the optimization of the entire system, multi-stage rectification should be used to distribute the dehydration load, that is, part of the dehydration task is completed in the gas separation section of the reactor, and part of it is completed in the subsequent tail gas recovery unit. The number of theoretical plates in the suitable separation section of the reactor given by the present invention is 5-15 theoretical plates, and the height of the corresponding gas separation section accounts for 10-30% of the total tower height.

Between the bottom plate of the gas separation section and the liquid level of the three-phase reaction section is the free space of the reactor, whose height accounts for 5-15% of the total tower height, mainly due to the fluctuation and splashing of the liquid level Leave room to prevent it from affecting the gas separation section.

3. Reactor aspect ratio

The aspect ratio of the reactor (that is, the ratio of height to diameter) is an important design parameter. When the reactor capacity and volume are constant, different aspect ratios correspond to different reactor diameters and heights, resulting in superficial gas velocity and downside pressures vary. The determination of the height-to-diameter ratio needs to take into account the flow and mass transfer requirements of the gas separation section, the gas-liquid mass transfer and radial distribution of gas velocity in the three-phase reaction section, gas holdup, and reactor residence time. The suitable height-to-diameter ratio of the three-phase reaction section given by the present invention is between 3 and 7, and the corresponding heights of the gas separation section and the free space section can be calculated according to this parameter.

4. Oxidation reaction conditions

Reaction temperature is the most important process design parameter. According to the principle of chemical engineering, adopting a higher reaction temperature can increase the energy grade of the by-product steam, which is beneficial to energy saving; high temperature reaction conditions are also conducive to the growth of TA crystals, which is convenient for filtration and drying; at the same time, the amount of catalyst and the amount of impurities generated at high temperature are also reduced. Less, which is beneficial to the purification of mother liquor. However, too high temperature will intensify the combustion side reaction of raw materials and solvents, increase material loss, and increase equipment investment. Considering the above advantages and disadvantages comprehensively, according to the requirement of minimizing energy consumption and material consumption of the whole system, the suitable reaction temperature given by the present invention is 185-195° C., and the corresponding pressure is 1.0-1.8 MPa.

The reactor residence time (defined as the liquid phase volume in the reactor/outlet liquid volume flow rate) mainly affects the depth of oxidation reaction, crystal particle size and reactor capacity. The suitable reactor residence time given by the present invention is 50-120 minutes.

Other reaction conditions are:

Feed solvent ratio—that is, the mass ratio of the raw material p-xylene fed into the reactor to the aqueous acetic acid solvent is 1:3 to 1:10.

Catalyst concentration and ratio - the total concentration of cobalt, manganese and bromine in the liquid phase in the reactor is 1000-3000ppm, wherein the atomic ratio of cobalt/manganese is 3:1-1:3, and the metal catalyst (cobalt The atomic ratio of + manganese) to bromine is 0.5 to 2.0:1.

Water content—the water content of the liquid phase in the reactor is 6-15% (w/w).

Tail gas oxygen concentration - the volume ratio of oxygen to nitrogen in the tail gas is 3-6%.

2. Slurry processing unit

The solid terephthalic acid generated in the bubble column reactor is discharged from the slurry pipe at the bottom of the reactor together with the mother liquor, and enters the slurry treatment unit, as shown in Figure 3. The slurry treatment unit includes multi-stage evaporators 301, 302, 303, filter 304, dryer 305, mother liquor purification device 306, and related auxiliary equipment. Its function is to carry out liquid-solid separation to recover terephthalic acid, recover the energy of the slurry stream, and purify the mother liquor to control the impurity content in the liquid phase. They are described as follows:

1. Multi-stage evaporator

The temperature and pressure of the slurry discharged from the bubble column reactor are relatively high. In order to facilitate liquid-solid separation and crystallize a small part of TA dissolved in the liquid phase, it is necessary to reduce the temperature of the slurry under reduced pressure. Although the purpose can also be achieved by using the first-stage flash method, it is difficult to recover the energy in the slurry. In addition, some reaction intermediates in the reaction mother liquor, such as p-toluic acid and p-tolualdehyde, have not been fully converted into terephthalic acid, and additional oxidation is required. Therefore, setting up a multi-stage evaporator can recover steam energy of different energy levels on the one hand, and on the other hand, can also pass some oxygen-containing gas into the evaporator for supplementary oxidation to increase the yield of TA.

The multi-stage evaporators 301-303 are composed of three stirred tanks connected in series, and each tank has the same volume and structure. The temperature of the first evaporator 301 is 2-5°C lower than that of the bubble column reactor, and the pressure is 0.1-0.3MP lower so that the slurry is automatically entered into the evaporation tank from the slurry pipe under the pressure difference. The residence time of the slurry in the first evaporator is 20-40 minutes. Oxygen-containing gas is introduced from the bottom of the first evaporator to promote the deep oxidation reaction. The total amount of gas introduced accounts for 2-8% of the gas input to the bubble column reactor. The O ₂ /N ₂ in the tail gas Ratio (volume ratio) is controlled at 3～6%. The evaporated steam and reaction tail gas flow out from the top of the tank after depressurization, and the condensate recovered after being condensed by a heat exchanger is introduced into the acetic acid storage tank, and the non-condensable tail gas is introduced into the tail gas treatment unit for absorption and purification, and the by-product of the heat exchanger Steam is used to provide heat or power for other equipment. The temperature and pressure of the second evaporator 302 are further lowered to promote more solvent evaporation, and the steam is directly introduced into the lower part of the rectification tower G16 of the tail gas treatment unit for dehydration, and also provides part of the heat source for the rectification tower. The last stage evaporator 303 reduces the pressure to normal pressure or negative pressure for slurry filtration.

2. Filtration, drying and mother liquor purification

The slurry coming out of the multi-stage slurry evaporator enters the filter 304 to filter and separate the mother liquor and solid, and at the same time pass a certain amount of acetic acid for washing to reduce the impurities and catalyst content in the solid. The filtered solids are sent to the drier 305 for drying to recover acetic acid. Most of the filtered mother liquor is returned to the bubble column reactor, and a small part is pumped out to the mother liquor purification device 306 for purification treatment to control the impurity content in the circulating mother liquor and recover the metal catalyst components in the mother liquor.

The purification of the mother liquor adopts the method of evaporation and concentration to recover the solvent acetic acid, and the obtained impurity residue is dissolved and extracted with water to recover the catalyst, and the catalyst can also be recovered by the oxalic acid precipitation method.

3. Exhaust gas recovery unit

The gas flowing out from the top of the bubble column reactor contains a large amount of acetic acid and water vapor, some non-condensable gases (N ₂ , CO ₂ , CO), and a small amount of other organic components (p-xylene, methyl bromide, methyl acetate etc.), need to be processed through the tail gas recovery unit, recover energy and solvent acetic acid, perform tail gas purification, and then discharge. The flow structure of the tail gas treatment unit is also shown in Figure 3, including multistage condensers 205, 206, 207, 208, azeotropic rectification column groups 216, 219, 220, tail gas absorption and purification devices 213, 214, and steam turbines and Expansion machines and other equipment are described as follows:

1. Multi-stage condenser

The multi-stage condensers 205-208 at the top of the bubble column reactor are also steam generators, consisting of several shell-and-tube heat exchangers connected in series, in which the reaction tail gas passes through the tube side, and the cooling water and by-product steam pass through the shell side . On the one hand, the function of the condenser is to condense the reaction tail gas, condense the solvent acetic acid and return it to the reactor, and on the other hand, recover the heat of reaction by producing steam. Multi-stage condensers are generally set with 3 to 4 stages to generate steam with different energy grades. A part of the condensate flows back into the reactor from the top of the tower, and a part is drawn out for dehydration to remove the water generated by the reaction and control the water content in the reactor. The high-pressure steam generated by the first condenser and the medium-pressure steam generated by the third condenser are mainly used to drive the steam turbine, and the medium-pressure steam generated by the second-stage condenser is mainly used to provide the reboiler heat source for the azeotropic distillation column . The condensers after the third stage use cooling water to condense the tail gas to a given temperature, without by-product steam.

2. Distillation column group

A part of the condensate after tail gas condensation needs to be taken out for dehydration to control the water content in the reactor. The extracted liquid is introduced into rectification column 216 to separate acetic acid and water. The rectification tower provided by the present invention adopts the azeotropic rectification method and multi-strand feed measures to facilitate energy saving. The condensate (containing water and acetic acid) to be separated is fed from the middle of the tower, another stream of water-containing acetic acid vapor from the second evaporator of the slurry treatment unit is fed from the lower part of the tower, and the separated acetic acid flows out from the bottom of the tower, and part of it is refluxed to the drum Bubble column reactor, part of which is provided to other equipment as washing liquid. The water vapor flows out from the top of the tower, part of it is drawn out after condensation, and part of it returns to the top of the tower. The reflux ratio is generally maintained at 1.5 to 4.5 to control the concentration of acetic acid in the tail gas at the top of the tower. The heat source of the reboiler partly comes from the by-produced steam of the second condenser at the top of the bubble column reactor, and partly comes from the water-containing acetic acid vapor sent from the second slurry evaporator. Entrainer can choose n-butyl acetate, or n-propyl acetate. The temperature of the reboiler at the bottom of the tower is 110-130°C, and the temperature of the tail gas at the top of the tower is 80-100°C.

In addition to water, the condensate at the top of the rectification tower is also rich in entrainer, part of the condensate is pumped into the oil-water separator 218 for stratification, and the upper oil phase is introduced into the entrainer recovery tower 219 for rectification and dehydration. In addition, the feed to the dehydration tower 216 also contains a small amount of p-xylene, which is easily enriched in the entrainer during rectification, so a p-xylene recovery tower 220 is also required to purify the azeotrope. In this way, the rectification tower bearing the solvent dehydration load is composed of three towers: an azeotropic rectification tower 216 , an entrainer recovery tower 219 and a p-xylene recovery tower 220 .

3. Exhaust gas absorption and purification device

After multi-stage condensation, the tail gas from the reactor still contains a small amount of acetic acid, p-xylene, and other organic components, which should be further recovered and purified. The tail gas first absorbs p-xylene and acetic acid through the absorption tower 213, and the absorbed tail gas enters the purification device 214 for catalytic oxidation or combustion. On the one hand, organic substances are removed, and on the other hand, the temperature of the tail gas is increased. The purified exhaust gas enters the exhaust gas expander to do work, recover the expansion energy of the gas, and then discharge it. The compressor 202, the expander 204, the steam turbine 203 and the starter motor 201 are connected in a coaxial manner to save energy.

4. System energy integration and process optimization

The oxidation of p-xylene is a strong exothermic reaction, and the effective use of reaction heat is the main measure for process energy saving. Most of the oxidation reaction heat is recovered through the tail gas condensers at all levels of the reactor to generate steam of different energy levels, part of which is used to drive the steam turbine and provide power for the compressor, and part of it is used to provide power for the dehydration tower, dryer, mother liquor purification device, etc. The device provides the heat source. Another part of the energy is recovered by the steam generated by the multi-stage slurry evaporator. According to the system engineering method, the present invention comprehensively considers the supply-demand balance between the energy generated by the reaction and the energy used for separation by each unit equipment. The distribution of steam and the heat exchange between streams are optimized according to the system energy integration method, so that the heating and cooling load of the required public works of the system is minimized. According to the process design and typical reaction conditions shown in Figure 3, the energy supply and energy demand load of each unit stream of the system are listed in Figure 4, where the steam and The tail gas is a hot stream that provides energy, and the heat source required by the rectification tower group, dryer, and mother liquor purification device is a cold stream. The energy level matching and energy distribution between the streams are indicated by arrows in the figure. Among the hot streams on the left in the figure, the one with the largest production capacity and the highest energy level is the steam by-produced by the tail gas condensers at all levels. Apart from a part of the energy provided is used for heating the cold stream on the right, a considerable part is also used to drive the steam The turbine does work to power the compressor. The energy of doing work is indicated by the block diagram and arrows in the middle part of the figure. The heat stream with a temperature lower than 115°C, such as the steam at the top of the dehydration tower and the condensed reaction tail gas, has a low energy level and is no longer useful. This part of waste heat needs to be cooled by cooling water provided by the outside. The cold stream with a temperature higher than 135°C, such as the heat source required by the dryer, cannot be heated by the hot stream, and needs to be provided by external utilities, but this part of the energy demand is small. The dehydration tower consumes the most energy in the system except for the compressor (not shown). Its energy can be exchanged with the by-produced steam of the hot stream or directly matched with the hot stream.

The process conditions and energy matching schemes of each unit and equipment given in the present invention are determined according to the principle of minimum material consumption and energy consumption of the whole system, and relevant process parameters will be described through Example 1.

Description of drawings

Fig. 1 is a unit module and process task diagram of the terephthalic acid generation process;

Fig. 2 is the enlarged structural representation of bubble column reactor 100 provided by the present invention;

Fig. 3 is a schematic diagram of a process flow of the present invention;

Fig. 4 is the energy level and stream matching diagram of each unit stream in the production process of terephthalic acid.

Detailed ways

As shown in Figure 1: the production process includes three units: oxidation reaction unit, slurry treatment unit and gas treatment unit. The raw material p-xylene and oxygen-containing gas react in the bubble column reactor 100 to generate solid terephthalic acid; the reaction slurry enters the slurry processing unit, and the task and function of this unit is to perform liquid-solid separation and recovery of terephthalic acid 1. Purify the mother liquor to control the impurity content, recover part of the energy in the slurry, and send the obtained crude terephthalic acid into the refining process to further make purified terephthalic acid; the steam and tail gas from the bubble column reactor 100 enter Tail gas treatment unit, the task and function of this unit is to recover reaction heat and gas expansion energy, recover solvent acetic acid, and perform gas purification to meet emission standards.

As shown in Figure 2, the bubble column oxidation device for producing terephthalic acid comprises a bubble column reactor 100 uniform up and down, and the bubble column reactor 100 is followed by a gas treatment section 101 and a free space section 102 from top to bottom. And the three-phase reaction section 103. The gas treatment section 101 accounts for 10-30% of the total height of the bubble column reactor 100, the three-phase reaction section 103 accounts for 60-85% of the total height of the bubble column reactor 100, and the free space section 102 accounts for 5-15%. Packing or liquid-holding tray 108 is installed in the gas treatment section 101, and is used for reclaiming the solid particles and part of acetic acid entrained in the tail gas, and a gas distributor 107 is horizontally installed at the lower part of the three-phase reaction section 103, and the gas distributor 107 is connected with the drum. The air inlet 105 of the bubble tower is connected, and a raw material feed pipe 104 is provided at 1/4 to 1/2 of the height of the three-phase reaction section 103 above the gas distributor 107, a discharge pipe 106 is provided at the bottom of the tower, and a discharge pipe 106 is provided at the top of the tower. A tail gas pipe 110 and a condensate return pipe 109 are arranged, and the tail gas pipe 110 is connected with the inlet of the multi-stage condenser, and the liquid outlet of the condenser is connected with the condensate return pipe 109 and the subsequent rectifying tower 216 respectively, and the gas outlet of the condenser is connected with the The subsequent exhaust gas treatment unit is connected. Part of the condensate flows back into the bubble column reactor 100 through the return pipe 109, and the other part is sent to the subsequent rectification tower 216 for dehydration, and the non-condensable tail gas is sent to the follow-up tail gas treatment unit for further treatment.

As shown in Figure 3, the production process of terephthalic acid is implemented as follows: (301-306 are the main devices of the slurry processing unit, 201-221 are the main devices of the gas processing unit, and 1-48 are the main process pipelines and stream numbers )

Line 1 supplies the raw material p-xylene (PX) to the bubble column reactor, cobalt, manganese metal and bromide are added through line 2 as catalysts for the oxidation reaction, and at the same time, lines 3 and 5 provide the solvent acetic acid required for oxidation. The air compressor 202 compresses the normal pressure air to about 1.2-2.0MPa (about 0.2MPa higher than the pressure in the bubble column reactor R1), and passes it into the bubble column reactor 100 through the pipeline 4. The oxygen in the high pressure air as an oxidizing agent for the reaction. The reaction temperature is 185-195° C., and the pressure is 1.0-1.8 MPa. PX undergoes liquid-phase oxidation in the reactor to generate solid terephthalic acid (TA), and the residence time of the slurry is 50-120 minutes.

The steam and tail gas generated in the three-phase reaction section 103 enter the gas separation section 101 through the free space section 102, and come into countercurrent contact with the condensate returning from the top of the tower, and the solid terephthalic acid particles entrained in the gas are recovered and flow back to the lower part with the liquid In the three-phase reaction section 103, acetic acid and water are rectified and separated in the gas separation section 101, and the light component water is discharged from the tail gas pipe 110 along with the steam, and is multi-stage condensed through the pipeline 8 to the condenser, while the solvent acetic acid is concentrated , and return to the lower three-phase reaction section 103 together with the solid matter.

The tail gas flowing out from the top of the reactor enters the tube side of the multi-stage condenser 205-208 through the pipeline 8, and the shell side of the condenser takes cooling water and steam. The first-third-stage condensers 205-207 will generate saturated steam of different temperatures and pressures, which are respectively used as power for the steam turbine 203 and heat sources for other process equipment according to a certain ratio. About 2/3 of the steam generated by the first condenser 205 is used to drive the steam turbine 203, and the rest is used for heating of other equipment; the steam generated by the second condenser 206 is mainly used to provide reboiling for the rectification column 216 The steam generated by the third condenser 207 is also mainly used to drive the steam turbine 203.

The temperature of the fluid at the outlet of the last-stage tail gas condenser 208 has dropped to 100-110°C, and it is sent to the gas-liquid separator 209, and most of the separated liquid phase is refluxed into the bubble column reactor 100 through the pipeline 7, A small part of the liquid is sent to the rectification column 216 through the pipeline 13 for dehydration of the solvent acetic acid. The gas separated by the vapor-liquid separator 209 is sent to the condenser 210 for further condensation and flash evaporation through the pipeline 10, and the liquid part is passed through the pipeline 12 to the rectifying tower 216 to separate acetic acid and water, and the gas is passed through the pipeline 11. into the tail gas absorption tower 213 bottom, reclaim acetic acid and p-xylene entrained therein.

In the tail gas absorption tower 213, the tail gas is in countercurrent contact with acetic acid and water respectively, and the p-xylene and acetic acid entrained in the tail gas are absorbed by the liquid phase acetic acid and water respectively. The acetic acid washing liquid comes from the acetic acid storage tank 215 and the condensing flash evaporator 212 respectively, and the washing liquid absorbs p-xylene and sends it back to the bubble column reactor 100 through the pipeline 23 . The tail gas after acid washing enters the upper part of the absorption tower, and the acetic acid in the gas phase is absorbed by water. The water washing liquid is sent from the entrainer recovery tower 219 through the pipeline 21, and is discharged from the side line after absorbing acetic acid, and is passed into the middle part of the rectification tower 216 through the pipeline 24 to separate acetic acid and water.

The tail gas discharged from the top of the absorption tower enters the tail gas purifier 214 for oxidation and temperature rise, and then is introduced into the tail gas expander 204 to recover expansion energy. The expander 204 is coaxial with the air compressor 202 , the steam turbine 203 and the starter motor 201 . The steam turbine 203 is driven by the steam by-produced by the tail gas condensers 205-207, converts the thermal energy of the steam into mechanical energy, and provides the energy required by the compressor together with the expander. The remaining energy of the steam is converted into electrical energy through the turbine, which can be transported to the outside. When the device is started, the motor 201 drives the air compressor to start the whole system to run. When the production process reaches a steady state, the motor drive is switched to the steam turbine and expander drive. The recovered energy has surplus power besides the self-sufficiency of the system. output.

Most of the tail gas condensate obtained from the gas-liquid separator 209 is refluxed to the bubble column reactor 100 through the pipeline 7, and a small part is drawn out for dehydration to remove the water generated by the reaction. The extracted liquid, together with the condensate in the condenser 210 and the washing liquid in the tail gas absorption tower 213, is sent to the middle of the rectification tower 216 for feeding through the pipeline 28, for rectification and dehydration. At the same time, the steam generated by the second evaporator 302 of the slurry processing unit is fed from the bottom of the rectification tower 216 to provide part of the energy required for dehydration, and the remaining energy is generated by the water vapor generated by the second condenser 206 through the reboiler heat exchanger supply. The entrainer (butyl acetate or propyl acetate) and water are added from the top of the tower through the pipeline 36, and the vapor phase product after rectification is rich in water and the entrainer, flows out from the top of the tower, and is condensed in the condenser 217 through the pipeline 35. In the oil-water separator 218, it is separated into a water phase and an oil phase. Part of the oil phase (water content 4-10%wt) is refluxed into the rectification column 216 through the pipeline 36 . The water phase (containing 1-5%wt of entrainer) is sent to the entrainer recovery tower 219 for rectification by pipeline 39, and the recovered entrainer (containing 45-60%wt of entrainer) is sent to oil-water phase through pipeline 38 The separator 218 separates the entrainer and water, so that the entrainer can be recycled. The oil phase of the oil-water separator 218 also contains a small amount of p-xylene, and a part of the oil phase is sent to the p-xylene recovery tower 220 by the pipeline 47, and its liquid phase product is rich in p-xylene, and is sent back to the bubble column by the pipeline 44 The gas phase product of the reactor 100 is rich in entrainer, which is sent back to the oil-water separator 218 through the pipeline 37 .

After the acetic acid is concentrated in the rectifying tower 216 (concentration>90%wt), it is discharged from the bottom and sent to the acetic acid storage tank 215 by the pipeline 29. A part of the recovered acetic acid is sent back to the bubble column reactor 100 through the pipeline 33, and a part is used as a cleaning liquid for slurry filtration and tail gas absorption.

The slurry discharged from the bottom of the bubble column reactor enters the first slurry evaporator 301 through the pipeline 6, where it is evaporated under reduced pressure on the one hand and undergoes supplementary oxidation on the other hand. The oxygen-containing gas is passed into the bottom of the slurry evaporator 301 through the pipeline 14, and the generated steam and tail gas enter the heat exchanger 211 through the pipeline 16, and are separated to obtain a vapor phase and a liquid phase. The liquid phase is directly sent to the acetic acid storage tank 215 through the pipeline 18, and the vapor phase is sent to the condensing flash tank 212 through the pipeline 17 for further separation. Also send to tail gas absorption tower 213 top recovery wherein acetic acid.

The slurry flowing out from the first slurry evaporator 301 is sent to the second slurry evaporator 302 through the pipeline 15 for further decompression evaporation, and the generated steam is sent to the bottom of the rectification column 216 through the pipeline 30 to provide part of the energy required for dehydration , the concentrated slurry is directly transported to the third slurry evaporator 303 through the pipeline 31, and the pressure of the evaporator is reduced to normal pressure, and the generated steam is condensed by the condenser and then refluxed, so that the temperature is reduced to 85-95°C, so that Subsequent filtration and drying.

The slurry flowing out from the third slurry evaporator 303 is sent to the filter 304 through the pipeline 32 to separate the mother liquor and solids, and at the same time, the acetic acid tank 215 sends part of the acetic acid to the filter 304 through the pipeline 34 to wash the solids. Most of the mother liquor in the filter 304 flows back to the bubble column reactor 100 through the pipeline 42, and a small part (5-10%) is extracted for mother liquor purification and catalyst recovery, and the recovered acetic acid is sent to the acetic acid storage tank 215 through the pipeline 27. The filtered and washed solid is sent to the dryer 305 through the pipeline 40 for drying, so that the moisture content is reduced from 10-15% to 0.05-1%. The dried crude TA will be sent to the refining unit for further hydrogenation treatment to obtain purified terephthalic acid.

As shown in Figure 4, the energy of each stream is arranged in order from top to bottom according to the temperature. The left part is the heat stream that provides energy, its length indicates the amount of energy that can be provided, and its temperature indicates the level of energy, which is a measure of effective energy. The right part is the cold stream that needs to be heated, its length indicates the energy consumption, and the temperature indicates the required energy level. The arrows in the figure indicate the energy distribution or matching between the specified streams, the numbers on the arrows indicate the amount of energy provided by each stream, and the block diagram and arrows in the middle part indicate the conversion of thermal energy into mechanical energy to do work.

The advantages and implementation details of the present invention will be further illustrated below through Example 1 and Comparative Examples.

Example 1.

Adopt the method and equipment provided by the present invention to carry out the terephthalic acid process and device design with an annual output of 600,000 tons, the annual production time is 7600 hours, and the functions and tasks of each unit determined by Fig. 1 are obtained through system combination optimization as shown in Fig. 3 The process conditions of each unit equipment are determined according to the principle of minimizing material consumption and energy consumption, and the energy matching between each stream is optimized according to the principle of system energy integration. The obtained relevant process parameters are listed as follows according to the unit order:

1. Bubble column reactor

The two bubble tower reactors shown in Figure 2 are operated in parallel, the production capacity of a single reactor is 300,000 tons/year, and a single set of equipment is used for slurry treatment and tail gas treatment. The reactor size and process parameters are listed in Table 1, and the reaction results and related indicators are listed in Table 2.

                                                                                                                                               

                                   

Diameter Total height Reaction section body Reaction section temperature PX throughput

(MPa, absolute (ppmHAc)

(m) degree (m) product (m ³ ) (℃) (10 ³ kg/h)

Co Mn Br

4.70 32 360 190 1.24 51.1 380 380 760

Exhaust gas separation

Feed HAc/PX Air flow rate Condensate return Condensate return

Section Theory Tower

Mass ratio (standard m ³ /h) Ratio Temperature (℃)

Number of boards

4:1 9.9×10 ⁴ 22 102 7

                                                                     

PX conversion rate TA yield (%) gas holdup in reaction section solid holdup in reaction section TA crystal particle size water content in mother liquor

(%) (%) (%) (%) (um) (mass%)

99.5 96 46 27.8 87 9.2

Mother liquor 4-CBA Solid phase 4-CBA Tail gas CO _X concentration Tail gas acetic acid concentration Tail gas PX concentration Tail gas oxygen concentration

Concentration (ppm) Concentration (ppm) Degree (CO _X /N ₂ %) (HAC/N ₂ %) (PX/N ₂ %) (O ₂ /N ₂ %)

2584 4773 1.67 0.93 7.7× ^10-5 3.5

In Table 1 and Table 2, PX refers to p-xylene, TA refers to terephthalic acid, CO _X concentration refers to the sum of CO and CO concentrations, 4-CBA refers to p-carboxybenzaldehyde (the _same in the following tables), HAC refers to acetic acid. Mother liquor 4-CBA concentration = liquid phase 4-CBA mass/solvent acetic acid and water mass, solid phase 4-CBA concentration = solid phase 4-CBA mass/solid phase TA mass.

2. Slurry processing unit

Three-stage slurry evaporators are used to decompress and evaporate step by step. The first slurry evaporator is fed with a small amount of air for deep oxidation, the steam generated by the second slurry evaporator is directly passed into the dehydration tower, and the slurry at the outlet of the third evaporator enters the filter. Machine and drying machine to produce crude TA products. The technological parameters of the three-stage evaporator and filter drying equipment are listed in Tables 3-6.

Table 3 Process parameters of the first slurry evaporator

Slurry 4-CBA Concentration

volume pressure temperature temperature solid holdup residence time air flow rate

(ppm)

(m ³ ) (MPa) (℃) (%) (min) (Nm ³ /h)

Solid phase Liquid phase

87.2 1.095 185 32 33 3.1×10 ³ 2300 <100

Table 4 Process parameters of the second slurry evaporator

volume pressure temperature temperature solid holdup residence time steam production rate

(m ³ ) (MPa) (°C) (%) (min) (kg/h)

66 0.217 136 41 34.6 52996

Table 5 Process parameters of the third slurry evaporator

volume pressure temperature temperature solid holdup water content of mother liquor residence time

(m ³ ) (MPa) (℃) (%) Amount (%) (min)

66 0.05 96 41 5.9 34.6

Table 6 Filter drying process parameters

Filter Dryer Filtration area Area load Moisture content of filter cake (m ² ) (kgTA/m ² .h) Rate (%) Heat exchange area Heat load Discharge moisture Catalyst residue (m ² ) (KW) Rate (%) Retention (ppm) 38.7 2.2×10 ³ 14.7 4500 2.8×10 ³ 0.5 35.4

3. Exhaust gas treatment unit

The tail gas at the top of the bubble column reactor is condensed in four stages, the first three stages by-product steam, the fourth stage is condensed with condensed water to 102°C, 95% of the condensate is refluxed to the reactor tail gas treatment section, and 5% is drawn out and led to the dehydration tower Distillation separation. The heat transfer load and operating conditions of the four-stage condenser are listed in Table 7.

                                                                                       

First condenser Second condenser Exhaust gas outlet By-product steam Heat exchange area Heat exchange load temperature (°C) Temperature (°C) (m ² ) (KW) Exhaust gas outlet By-product steam Heat exchange area Heat exchange load temperature (°C) Temperature (°C) (m ² ) (KW) 165 140 3200 6.5×10 ⁴ 155 130 1900 3.3×10 ⁴

Third Condenser Fourth Condenser Exhaust gas outlet By-product steam Heat exchange area Heat exchange load temperature (°C) Temperature (°C) (m ² ) (KW) Exhaust gas outlet Heat exchange area Heat exchange load temperature (°C) (m ² ) (KW) 130 115 2300 4.4×10 ⁴ 102 2000 2.8×10 ⁴

The tail gas condensate and the steam of the second slurry evaporator are dehydrated by azeotropic rectification, the entrainer is n-butyl acetate (NBA), the condensate enters from the middle of the tower, and the steam enters from the bottom of the tower. The heat source of the reboiler is the steam by-produced by the second condenser of the reactor tail gas. The process parameters of the rectification column group are given in Tables 8-10.

                                                                                                                      

Theoretical Plates Liquid Phase Feed Vapor Phase Feed Column Top Pressure Column Bottom Pressure Reflux Ratio

Number Location Location (MPa) (MPa)

30 15 30 0.1 0.12 2.1

Reboiler Negative Condenser Negative Tower Top Acid Concentration Tower Bottom Acid Concentration Tower Top Temperature Tower Bottom Temperature

Charge (KW) Charge (KW) Degree (mass%) Degree (mass%) (°C) (°C)

1.7×10 ⁴ 2.2×10 ⁴ <0.1 92 92 119

                                                                                           

Theoretical column Feed level Middle outlet Reboiler load Bottom outlet Total Middle outlet Total Top outlet

Reflux ratio

Number of Plates Placement Material Position (KW) Boiling Agent Concentration Boiling Agent Concentration Boiling Agent Concentration

21 12 10 20 3.0×10 ³ <0.1 86 <0.1

                                                                                                                                                                                                                                                 Table 10 Process parameters of p-xylene recovery tower

Reboiler Duty Bottom Discharge Top Discharge

Theory Feed Position Reflux Ratio

(KW) PX Concentration (%) PX Concentration (%)

20 2 1.3 5.0×10 ² 99 3.5

After four-stage condensation, the tail gas of the reactor enters the tail gas absorption tower to absorb p-xylene and acetic acid. The absorption tower adopts a two-stage structure. The lower part first absorbs the p-xylene contained in the tail gas with acetic acid solution, and the upper part absorbs the acetic acid contained in the tail gas with water. The relevant process parameters are given in Table 11.

Table 11 Process parameters of exhaust absorption tower

Spraying in the lower section Concentration of gas components at the outlet of the lower section Concentration of gas components at the outlet of the upper section

Upper stage spraying

Acetic acid flow rate

                                                                                                           

(kg/h)

2.7×10 ⁴ 2.4 0.3 1.4×10 ^-4 4.2×10 ³ 3.3 2.6×10 ^-4 1.4×10 ^-4

4. Main material consumption and energy consumption indicators

According to the above embodiment, table 12 has provided the material consumption and energy consumption per ton of TA in the process, wherein the consumption of acetic acid and p-xylene has only considered the main device bubble column reactor, dehydration tower, tail gas absorption tower, drying The material consumption of machines, etc., does not include the material consumption of other auxiliary equipment. Energy consumption only considers the energy consumption of the main equipment in Figure 3, excluding the power consumption of auxiliary equipment, instruments, etc. The recovered steam energy is also uniformly converted into electrical energy for consideration, so the power consumption of main equipment given in the table is the net consumption of the system calculated in terms of electrical energy after energy recovery is considered.

Table 12 Material Consumption and Energy Consumption

     Main equipment material consumption (kg/tTA)                                                                          

p-xylene acetic acid entrainer catalyst (KW.h/tTA) (t/tTA) consumption (t/tTA)

648.5 38.5 0.5 0.2 41.5 24.4 0.05

Comparative example 1

In order to further illustrate the advantages of the present invention, the production process provided by the present invention can be compared with existing typical processes. As comparative example 1, the flow process of reference comparison is to adopt conventional stirred tank reactor to replace the bubble column reactor of the present invention, and the dehydration tower of the tail gas treatment unit is to adopt common rectification to replace the common rectification that the present invention adopts simultaneously The configuration and conditions of boiling rectification, slurry treatment unit and other equipment are the same as those of the present invention.

Tables 13 to 15 give the main equipment and process parameters of the comparison process with an annual output of 600,000 tons of terephthalic acid, and Table 16 lists the relevant material consumption and energy consumption indicators.

Table 13 Stirred tank bubble column reactor process parameters

Catalyst Concentration

Reactor platform Single reaction Reaction section temperature Tower top pressure PX handling capacity Feed HAc/

(ppmHAc)

Segment volume (m ³ ) degree (°C) (MPa, absolute pressure) (10 ³ kg/h) PX mass ratio

Co Mn Br

2 286.0 190 1300.0 51.1 4:1 480 480 960

                                                                                                       

PX conversion rate TA yield (%) gas content in reaction section solid content in reaction section TA crystal particle size water content in mother liquor

(%) Rate (%) Rate (%) (um) (Quality%)

99.5 96 38 31 103.0 9.2

Mother liquor 4-CBA Solid phase 4-CBA Condensate extraction Tail gas oxygen concentration Tail gas CO _X concentration Tail gas PX concentration

Concentration (ppm) Concentration (ppm) Flow rate (kg/h) (volume %) Degree (volume %) (volume %)

2300 4500 734600 3.5 1.7 5.2×10 ^-4

Table 15 Process parameters of ordinary distillation dehydration tower

                                                           

Number of theoretical plates

Position (MPa) (MPa) (°C)

40 25 40 0.1 0.16 98

Bottom Temperature Reboiler Duty Condenser Duty Top Acid Concentration Column Bottom Acid Concentration

Reflux ratio

(°C) (KW) (KW) (Mass%) (Mass%)

126 15 3.4×10 ⁴ 3.9×10 ⁴ <0.1 95

                                                    Table 16 compares the material consumption and energy consumption of main equipment in process 1

 Material consumption of main equipment (kg/tTA)  Power consumption of main equipment  Consumption of cooling water  Consumption of externally supplied steam

P-Xylene Acetic Acid Catalyst (KW.h/tTA) (t/tTA) Consumption (t/tTA)

648.5 38.5 0.2 62.3 26.7 0.05

Comparing the process conditions and consumption index of the above-mentioned comparison process with the relevant results of Example 1, it can be seen that the production of terephthalic acid can also be realized by using the comparison process, but the energy consumption is higher than the index given in Example 1. The reasons are mainly reflected in two aspects: one is that the bubble column reactor used in the comparison process is a current general stirred tank reactor, and there is no gas separation section, so that the heat of reaction cannot be used for solvent separation more directly and effectively; Ordinary distillation is used for solvent dehydration, which consumes more energy than azeotropic distillation. In addition, the stirring power consumption of the reactor also slightly increases the energy consumption. Therefore, the above method and process are obviously inferior to the method and process provided by the present invention in terms of energy saving.

Comparative example 2

The difference between the flow process of the second reference comparison and the present invention lies in the slurry treatment unit, the comparative flow process adopts the one-stage flash method to concentrate the slurry at the outlet of the bubble column reactor, directly reduces the pressure to normal pressure, and then passes through two-stage centrifugation Machine separation of mother liquor and solids, to replace the multi-stage evaporation and filtration methods used in the present invention. The design of the bubble column reactor and the tail gas treatment unit is the same as that of the present invention. For the above comparison process, under the same 600,000 tons/year production capacity and oxidation reaction conditions, the relevant material consumption and energy consumption indicators are listed in Table 17.

          Table 17 compares the main equipment material consumption and energy consumption of process 2

     Main equipment material consumption (kg/tTA)                                                                                                  

p-xylene acetic acid catalyst entrainer (KW.h/tTA) (t/tTA) consumption (t/tTA)

648.0 37.7 0.2 0.5 73.1 26.2 0.05

Compared with the process given in this patent, the material consumption of comparative process 2 is slightly lower, mainly because the supplementary oxidation consumption of the first slurry evaporator is omitted. However, the energy consumption of this process is the largest among the above three processes. The main reason is that the energy in the slurry has not been recycled after the slurry treatment system adopts the first-stage flash method. At the same time, the TA particles broken by the centrifuge increase the moisture content of the solid, which also increases the drying load and energy consumption. Therefore, the method that comparative example 2 provides is also not as good as the method that the present invention provides in terms of energy saving

Finally, it should also be noted that what is listed above is only a specific embodiment of the present invention. Obviously, the present invention is not limited to the above embodiments, and many variations are possible. All deformations that can be directly derived or associated by those skilled in the art from the content disclosed in the present invention should be considered as the protection scope of the present invention.

Claims (8)

1, a kind of terephthalic acid production plant comprises bubbling column reactor (100), slurry treatment unit, exhaust gas processing device, it is characterized in that:

Described bubbling column reactor (100) is straight tubular, inside is followed successively by gas processing section (101) from top to bottom, free space segment (102) and three phase reaction section (103), described gas processing section (101) accounts for 10～30% of bubbling column reactor (100) height overall, and three phase reaction section (103) accounts for 60～85% of bubbling column reactor (100) height overall; Filler or supporting liquid column plate (108) are installed in gas processing section (101), lower horizontal at three phase reaction section (103) is equipped with gas distributor (107), gas distributor (107) links to each other with bubble tower inlet mouth (105), on the three phase reaction section (103) of gas distributor (107) top, be provided with material feeding tube (104), be provided with discharge nozzle (106) at the bottom of the tower, cat head is provided with offgas duct (110) and phlegma return line (109); Described bubble tower inlet mouth (105) links to each other with air compressor pump (202);

Described slurry treatment unit comprises the multi-step evaporator that links to each other with discharge nozzle (106), the upper outlet of multi-step evaporator by heat exchanger (211) back, link to each other with acetic acid tank (215), tail gas absorber (213) respectively, rectifying tower (216) is continuous, the following outlet of multi-step evaporator links to each other with filter (304), and the outlet of filter (304) links to each other with material feeding tube (104), mother liquor purification device (306), drying machine (305) respectively; The outlet of mother liquor purification device (306) links to each other with acetic acid tank (215);

Described exhaust gas processing device comprises the multi-stage condensing device, the shell side of the byproduct steam of multi-stage condensing device, link to each other with steam turbine (203) thermal source required with processing unit is provided respectively, the tube side of described multi-stage condensing device links to each other with phlegma return line (109), rectifying tower (216), tail gas absorber (213) respectively after through vapour liquid separator (209); Rectifying tower (216) links to each other with the outlet of heat exchanger (211), the following outlet of rectifying tower (216) links to each other with material feeding tube (104), acetic acid tank (215) respectively, the upper outlet of rectifying tower (216) links to each other with phlegma return line (109), water-and-oil separator (218) respectively, and the outlet of water-and-oil separator (218) links to each other with entrainer recovery tower (219), p-Xylol recovery tower (220) respectively; The outlet of tail gas absorber (213) links to each other with exhaust purifier (214).

2, a kind of terephthalic acid production plant according to claim 1 is characterized in that: the theoretical plate number of described gas processing section (101) is 5～15, and the height of oxidation reactor three phase reaction section (103) and diameter ratio are 3～7.

3, a kind of terephthalic acid production plant according to claim 1 and 2 is characterized in that: the outlet of described exhaust purifier (214) links to each other with tail gas decompressor (204).

4, a kind of terephthalic acid production plant according to claim 3, it is characterized in that: described multi-step evaporator comprises first vaporizer (301), second vaporizer (302), the 3rd vaporizer (303) of polyphone successively, described first vaporizer (301) is provided with feeding oxygen tube (14), and described feeding oxygen tube (14) links to each other with air compressor pump (202).

5, a kind of terephthalic acid production plant according to claim 4 is characterized in that: described multi-stage condensing device comprises first condenser (205), second condenser (206), the 3rd condenser (207), the 4th condenser (208) of polyphone successively; Electric motor (201) and described air compressor pump (202), steam turbine (203), the coaxial connection of tail gas decompressor (204).

6, a kind of terephthalic acid production method is characterized in that, comprises the steps:

1), oxidizing reaction: behind bubble tower inlet mouth (105) process gas distributor (107), the three phase reaction section (103) that enters respectively in the bubbling column reactor (100) carries out oxidizing reaction from material feeding tube (104), oxygen-containing gas for p-Xylol and solvent acetic acid; Reaction gained gas is through gas processing section (101) back, from offgas duct (110) discharge, and the slurry of reaction gained is discharged from discharge nozzle (106);

2), slurry is handled: the slurry of discharging from discharge nozzle (106) cooling of reducing pressure in multi-step evaporator, enter heat exchanger (211) from the upper outlet expellant gas of multi-step evaporator and carry out heat exchange, the phlegma of gained is introduced acetic acid tank (215), non-condensable gases is introduced tail gas absorber (213) and is handled, and steam introducing rectifying tower (216) dewaters and part of heat energy is provided; The slurry of discharging from the following outlet of multi-step evaporator is divided into mother liquor and solid through the filtration of filter (304), solid enters in the drying machine (305) dry to reclaim acetic acid, most mother liquor is got back to three phase reaction section (103) through material feeding tube (104), the mother liquor of small part carries out purifying treatment through mother liquor purification device (306), reclaim the metal catalyst composition in the mother liquor simultaneously, the mother liquor after the purification enters acetic acid tank (215);

3), vent gas treatment: after the processing of tail gas through 3～4 grades of condensers of offgas duct (110) discharge, the steam of generation drives steam turbine (203) and thermal source is provided for rectifying tower (216) by shell side, and the water coolant of generation is by shell side condensation tail gas; Described condensation tail gas from the tube side of multi-stage condensing device come out, through after vapour liquid separator (209), tail gas purifies by tail gas absorber (213), exhaust purifier (214) successively, a phlegma part turns back to inner, a part of rectifying tower (216) separating acetic acid and water, also introducing rectifying tower (216) separating acetic acid and the water that heat exchanger (211) comes out introduced of bubbling column reactor (100) by phlegma return line (109); Acetic acid after the separation partly enters bubbling column reactor (100) inside by material feeding tube (104), and part enters acetic acid tank (215); The phlegma of the gas condensation gained that comes out from rectifying tower (216) upper outlet, part enters bubbling column reactor (100) inside through return line (109), the part extraction enters water-and-oil separator (218) and separates, separating obtained oil phase is introduced entrainer recovery tower entrainer recovery tower (219) again and is carried out rectifying and dewatering, and entrainer enters p-Xylol recovery tower (220) to purify entrainer.

7, a kind of terephthalic acid production method according to claim 6, it is characterized in that: described three phase reaction section (103) interior reaction temperature is 185～195 ℃, pressure 1.0～1.8MP, catalyst cobalt, manganese, three kinds of ionic total mass concentrations of bromine are 1000～3000ppm, the atomic ratio of cobalt manganese is 3.0～0.3, and the atomic ratio of metal catalyst ion and bromide anion (ratio) is 0.5～2.5.

8, a kind of terephthalic acid production method according to claim 7 is characterized in that: slurry feeds oxygen-containing gas and carries out the supplemental reaction in first vaporizer (301).

Images