Intrinsic safety type efficient low-cost tubular reaction process and application

Abstract

The invention discloses an intrinsically safe, efficient, low-cost and tubular reaction process and application thereof. According to the technology, in the reaction process, energy generated by the reaction is comprehensively utilized to achieve balance, introduction of an external heat source is reduced, the production cost is reduced, the market competitiveness of products is improved, and meanwhile safety risks existing in the production process of an existing production technology are fundamentally eliminated. Through the design of the special'collision type 'structure mixer, the efficient conversion of kinetic energy of a two-phase material flow system is enhanced, the mixing effect is improved, and the reaction efficiency is enhanced. The process disclosed by the invention has a relatively good application prospect.

Description

Technical Field

[0001] The present invention relates to a tubular chemical reaction process technology, belongs to the field of chemical reaction engineering, and specifically relates to a reaction process that can be used for nitration reaction, etherification reaction, oxidation reaction, esterification reaction, amination reaction, chlorination and diazotization reaction. Background Art

[0002] As a traditional industry in my country, the chemical industry holds a crucial position in the national economy. Chemical reaction processes are the heart of chemical production. With the continuous development and advancement of technology, continuous tubular reaction processes have gained widespread application in the chemical synthesis industry. The continuous improvement of people's living standards has led to a deeper pursuit of chemical synthesis. Consequently, the country has placed increasing emphasis on environmental protection and safe production. The development of clean, environmentally friendly, and safe chemical processes has become crucial, driving technical personnel to continuously optimize and improve chemical production processes, process control, and production management. This has led to a high demand for the development and application of intrinsically safe, continuous, and highly efficient production processes. To achieve these goals, technical personnel have developed numerous process technologies and have successfully implemented them. In particular, the application of tubular reaction and microchannel reaction technologies has enabled the optimization of these processes and improved safety.

[0003] Chinese patent publication CN102432410B discloses a method for producing nitro compounds by continuous tubular nitration. This method provides a novel continuous production process using tubular nitration and tubular washing. The apparatus used in the new process includes a tubular nitration reactor, tubular water and alkali washing equipment, a nitration mixture separator, and a water and alkali washing separator. Nitration feedstock and mixed acid flow, mix, react, and transfer heat within the tubular reactor. The separated nitrate undergoes tubular water and alkali washing to yield a qualified nitration product. The document discloses that the nitration reactor, primary water scrubber, alkali scrubber, secondary water scrubber, and nitric waste acid extraction reactor are tubular devices. The document also discloses the internal structural features of the tubular devices, including the use of empty tubes, internal chemical fillers, internal disturbance elements, or static mixers. The document further discloses that the tubular nitration reactor can be divided into multiple sections, and nitric acid can be added at the inlet of the tubular nitration reactor or separately at the inlet and other locations. The patent document discloses a tubular continuous nitration reaction method for producing nitro compounds, which significantly improves production efficiency compared to classic, traditional batch-type parallel and multi-stage series production methods. Compared to multi-stage parallel or series production methods, which have large inventory volumes and pose significant safety risks, pipeline reactors with smaller inventory volumes offer unparalleled advantages in production safety. The patent document also discloses that the composition of the mixed acid added through the tubular reactor can control the adiabatic temperature rise of the reactor, further improving safety. Pipeline reactions have undergone cross-level process upgrades and improvements to address industrial disadvantages such as improved mixing, mass and heat transfer, and low production efficiency. In actual production, the complex reaction process, especially the tendency for overreaction during the reaction process, such as the production of polynitro impurities, is not permitted. This patent document does not disclose other control measures or methods. Regarding energy utilization, particularly the reuse of waste acid, this patent document does not disclose much information.

[0004] Chinese patent publication CN105541667B discloses a method for continuous pipeline production of CLT acid nitrates. This method primarily utilizes a pipeline reactor, aided by a rational flow and temperature control system, to complete the sulfonation, chlorination, and nitration reactions within the pipeline. The patent also discloses the essential features of the pipeline reactor, which utilizes a commercially available SK-type pipeline liquid-liquid static mixer and / or SV-type pipeline liquid-liquid static mixer. The reaction process eliminates the need for intermediate purification, resulting in simple operation, high reaction efficiency, and high selectivity for the desired product. Compared to traditional batch reactors, the patented method for continuous pipeline production of CLT acid nitrates offers a smaller reactor volume, improved reaction stability, and higher product selectivity. However, the reaction process is prone to overreaction, such as overaddition of toluenesulfonic acid, resulting in polysulfonic acid-CLT acid sulfonates, or overaddition of nitration, resulting in polynitro-CLT acid nitrates. These impurities can carry over into the next process, compromising the quality of subsequent products and impacting their market competitiveness.

[0005] With the progress of modern material manufacturing technology and mechanical processing level, for chemical production with relatively fast reaction rate in nature, the promotion and application of microchannel reactor in the field of chemical synthesis is valued by technicians. Publication number is that the document patent of CN114478260B discloses a method for preparing m-nitrobenzaldehyde by microchannel reactor. The method provides a method for preparing m-nitrobenzaldehyde by microchannel reactor. This industrial preparation method is not only simple to operate, easy to control, and has high production efficiency (high yield, high purity), high safety, and is suitable for industrial continuous production. The specific microstructure that microchannel reactor has is the core component of the reactor, and its unique structure gives a series of high-quality performances. Microchannel reactor utilizes the ultra-high pressure drop of its two-phase fluid to complete the mixing of materials. The disadvantage brought by the advantage is that the power difference of the two-phase fluid cannot be formed, and energy cannot be fully transmitted. The complex flow channel in the reactor causes the internal microstructure (micrometer level) fluid law to be extremely difficult to control. As in the nitration reaction, the situation of excessive nitration is unavoidable, and the application occasion of the reaction is greatly restricted.

[0006] As a revolutionary technology, microchannel reactors involve research covering multiple disciplines such as microdynamics, microtribology, microfluidics, microheat transfer, microelectronics, and chemistry. Therefore, each component and microchannel of the microreactor must be precisely designed and researched, resulting in a high market price for microchannel reactors. The "micro" in microchannel reactor indicates that the channel size of the process fluid is at the micron level. Its complex structure also determines that the production capacity of a single microreactor cannot achieve the expected effect. The use of microchannel reactors is even more limited in the presence of solids or fluids that are prone to crystallization. The rigid demand for industrial production capacity requires the simultaneous use of multiple microchannel reactors, all of which has led to a continuous increase in investment costs for manufacturers.

[0007] Furthermore, prior art pipeline reactors often incorporate internal fillers or incorporate agitation elements. These spiral mixing elements create vortices and cross-flow within the pipeline, preventing immiscible or low-solubility two- or multi-phase fluids from stratifying and degrading, thereby improving reaction efficiency. However, due to factors such as pipeline resistance and liquid viscosity, the fluid's flow rate at the pipe wall decreases, leading to a decrease in the Reynolds coefficient and a sharp drop in the mixing efficiency of the two-phase fluids. This phenomenon is known as wall sticking. Prior art currently lacks effective technical solutions to prevent this phenomenon.

[0008] In view of the existing production process technology, when tubular reactions are applied to production processes such as nitration, etherification, oxidation, esterification, amination, chlorination and diazotization, the reaction process is prone to produce excessive impurities due to overreaction, thereby affecting the quality of the product; the energy generated during the reaction process cannot be well utilized, and due to factors such as pipeline resistance and liquid viscosity, the flow rate of the fluid on the pipe wall is reduced, the Reynolds coefficient decreases accordingly, and the mixing effect of the two-phase fluid is sharply reduced, resulting in wall sticking phenomenon. The present invention provides an intrinsically safe, efficient, low-cost tubular reaction process and application, which eliminates safety hazards in industrial production based on the intrinsic safety of the synthesis reaction. The process of the present invention can be flexibly adjusted in terms of process, equipment, automatic control, etc. according to the specific reaction characteristics in the reaction process to optimize the quality of product synthesis. In the reaction process and application process, the energy generated by the reaction itself is comprehensively utilized to achieve a balance, reduce the introduction of external heat sources, reduce production costs, and improve product market competitiveness. At the same time, the safety risks existing in the production process of existing production technology are fundamentally eliminated. The present invention uses a special "collision-type" structure mixer design to enhance the efficient conversion of kinetic energy of the two-phase logistics system, improve the mixing effect, and enhance the reaction efficiency. In order to achieve the above purpose, the present invention adopts the following technical solutions:

[0009] An intrinsically safe, efficient, low-cost, tubular reaction process comprising:

[0010] a) a raw material preparation unit, comprising a mixing kettle 110 for batching, a continuous phase feed pipeline 1110, a raw material pipeline 1111, a delivery pump 111, a delivery pipeline 1112 and a raw material tank 100;

[0011] b) a reaction unit, comprising a mixer 200, a pipeline reactor 210, and a separator 300;

[0012] c) an energy integration unit, comprising a flash evaporator 310, a distributor 320, and heat exchangers 330 and 340;

[0013] d) Post-processing unit, including a waste water tank 350, a waste acid tank 360 and a crude product tank 370.

[0014] The present invention also provides a reaction process and application, wherein the application is preferably a nitration reaction, an etherification reaction, an oxidation reaction, an esterification reaction, an amination reaction, a chlorination reaction and a diazotization reaction.

[0015] The present invention has the following beneficial effects:

[0016] 1. The process of the present invention is an intrinsically safe reaction process. Starting from its essence, it avoids the occurrence of energy loss in the industrial production process, which leads to a sharp increase in temperature and causes safety hazards.

[0017] 2. The process of the present invention uses a "contraction-diffusion-contraction" collision-type mixer, which fully utilizes the conversion of the energy difference between the two phases and has a good mass transfer effect. It simplifies the equipment, can operate stably over a wide output range, and optimizes operation.

[0018] 3. The process of the present invention optimizes the manufacturing level of the pipeline reactor, improves the operating conditions, and the unique internal hollow structure can replace the boundary layer of the pipe wall and part of the mainstream fluid, which has low resistance, eliminates dead zones, reduces the friction between the fluid and the pipe wall, and greatly eliminates the occurrence of wall sticking.

[0019] 4. The process of the present invention can select the form of mixer combination according to the requirements of reaction selectivity, and can use a single-stage or series-parallel multi-stage mixer. It can be adjusted within a wide range of product quality requirements and the product quality is controllable.

[0020] 5. The tubular reactor in the process of the present invention can carry out conventional heat transfer reactions and is also suitable for adiabatic reactions. The use of flash evaporation technology and reaction energy conversion balances the entire production process, significantly reducing the operating costs of the entire process.

[0021] 6. In the present invention, the ratio of the continuous phase or solvent is adjusted, and the volume of the circulating continuous phase is increased to ensure sufficient heat capacity and to keep the adiabatic temperature rise of the reaction within a safe production range. Continuous interface separation control and an automatic distributor can achieve rapid equilibrium between the continuous and dispersed phases. BRIEF DESCRIPTION OF THE DRAWINGS

[0022] Figure 1 Shows a schematic diagram of an intrinsically safe, efficient, low-cost, tubular reaction process

[0023] Figure 2 Schematic diagram of the process scheme of Example 1 is shown

[0024] Figure 3 Schematic diagram of the process scheme of Example 2 is shown

[0025] Figure 4 Schematic diagram of the process scheme of Example 3 is shown

[0026] Figure 5 Schematic diagram of the process scheme of Example 4 is shown

[0027] Figure 6 Schematic diagram showing the mixer structure

[0028] Figure 6 In the figure, there are continuous phase inlet ①, dispersed phase inlet ②, collision zone ③, diffusion zone ④, and throttling port ⑤.

[0029] Figure 7 Shown Figure 6 Schematic diagram of the mixing angle θ1 in the collision zone ③

[0030] Figure 8 Shown Figure 6 Schematic diagram of the cross-section angle θ2 of the throttle port ⑤

[0031] Figure 9 Shown Figure 6 Schematic diagram of the longitudinal length L of the diffusion zone ④

[0032] Figure 10 Shown Figure 9 Schematic diagram of the 5th cross section of the throttle and the 4th cross section of the diffusion area

[0033] Figure 10 In the figure, the cross-sectional area of ​​the throttle port ⑤ is S2, the cross-sectional area of ​​the diffusion zone ④ is S1, and the cross-sectional diameter of the diffusion zone ④ is R.

[0034] Figure 11 Shows the longitudinal and cross-sectional views of the blades in the mixing unit of the pipeline reactor

[0035] Figure 12 A 3D perspective view showing the blades of the mixing unit of a pipe reactor

[0036] In the accompanying drawings of the present invention, there are a mixing kettle 110, a continuous phase feed line 1110, a raw material line 1111, a delivery pump 111, a pump delivery line 1112, a raw material tank 100, a mixer 200, a mixer 201, a pipeline reactor 210, a separator 300, a separator light phase line 3002, a separator heavy phase line 3001, a flash evaporator 310, a flash evaporator gas phase line 3101, a distributor 320, heat exchangers 330 and 340, a heat exchanger outlet line 3300, a batching kettle return line 3403, a waste water tank 350, a waste acid tank 360, and a crude product tank 370.

[0037] in, Figure 4 A salt slurry tank 360 is shown.

[0038] in, Figure 4 and Figure 5 Flasher liquid phase line 3300 is shown. DETAILED DESCRIPTION

[0039] In order to make the purpose, technical solutions and advantages of the embodiments of the present invention clearer, the technical solutions in the embodiments of the present invention will be clearly and completely described below. Obviously, the described embodiments are part of the embodiments of the present invention, not all of the embodiments. Based on the embodiments of the present invention, all other embodiments obtained by ordinary technicians in this field without making creative work are within the scope of protection of the present invention. Unless otherwise expressly stated, throughout the specification and claims, the term "including" or its variations such as "comprising" or "including" will be understood to include the stated components or steps, without excluding other material components or steps.

[0040] Furthermore, in order to better illustrate the present invention, numerous specific details are given in the following detailed description.

[0041] Those skilled in the art will appreciate that the present invention can be practiced without certain specific details. In some embodiments, raw materials, methods, means, etc. well known to those skilled in the art are not described in detail in order to highlight the main purpose of the present invention.

[0042] The continuous phase described in the present invention can be recycled: for example, concentrated sulfuric acid in the nitration reaction is well known to be recycled as a catalyst in the nitration reaction, namely the so-called continuous phase; the other reaction raw material corresponding to the continuous phase is often defined as the dispersed phase.

[0043] The present invention first provides an intrinsically safe, efficient, low-cost, tubular reaction process, which comprises:

[0044] a) a raw material preparation unit, comprising a mixing kettle 110 for batching, a continuous phase feed pipeline 1110, a raw material pipeline 1111, a delivery pump 111, a delivery pipeline 1112 and a raw material tank 100;

[0045] b) a reaction unit, comprising a mixer 200, a pipeline reactor 210, and a separator 300;

[0046] c) an energy integration unit, comprising a flash evaporator 310, a distributor 320, and heat exchangers 330 and 340;

[0047] d) Post-processing unit, including a waste water tank 350, a waste acid tank 360 and a crude product tank 370.

[0048] Preferably, in the process described above, raw materials and a continuous phase are added to the mixing kettle 110 , and the raw materials are added from the top of the batching kettle 110 or from the inlet of the delivery pump 111 .

[0049] Preferably, the delivery pump 111 is a high-pressure turbine mixing pump.

[0050] Preferably, the mixer in the above process is a "converging-diverging-converging" collision-type structure, primarily comprising a continuous phase inlet, a dispersed phase inlet, a collision zone, a diffusion zone, and a throttle. The mixer can be used alone, in a two-stage cascade, or in a three-stage cascade.

[0051] The mixer of the present invention adopts a "contraction-diffusion-contraction" collisional structure. The continuous phase experiences its first energy release in the collision zone of the mixer. Due to the kinetic energy difference between the continuous and dispersed phases, the atomized continuous phase forms small droplets that collide with the dispersed phase at high speeds, creating a highly intertwined interface between the two phases and allowing them to dissolve in each other. This allows for efficient mixing of the two phases. The mixed fluid passes through the mixer's diffusion zone, where the channel area is enlarged and the two phases undergo a "diffusion" process. The diffusion zone can be configured in various shapes and lengths, allowing for a continuous reaction. After passing through the diffusion zone, the fluid reaches the throttle. By controlling the cross-sectional area of ​​the throttle throat, the two-phase mixture undergoes a second "contraction." The kinetic energy of the two-phase flow continues to be converted in the throttle throat, where the fluids continue to shear and mix, completing the second half of the reaction. By controlling the cross-sectional area of ​​the throttle throat, the flow rate of the two-phase mixture can reach subsonic levels. The entire process of the immiscible flows of the continuous and dispersed phases within the mixer completes a "contraction-diffusion-contraction" process, achieving energy transfer and thus a substance transformation process.

[0052] Preferably, the continuous phase enters from the continuous phase inlet at a high pressure of 0-25 bar, preferably 3-18 bar, and the raw materials enter from the dispersed phase inlet at a low pressure of 0-10 bar, preferably 1-6 bar.

[0053] The continuous phase mainly refers to the reaction system solvent or catalyst logistics system, and the dispersed phase mainly refers to the reaction raw materials. Compared with the dispersed phase, the continuous phase has a larger volume and has a significant impact on the diffusion, penetration, heat transfer, adhesion, swelling and other properties of the reaction logistics system. In order to further strengthen the role of the continuous phase in the logistics system, the continuous phase needs to enter at high pressure and exist in the form of small and dispersed particles. As the reaction raw material, the dispersed phase enters the mixer at low pressure, forming a kinetic energy potential difference, strengthening the interaction between the dispersed phase and the continuous phase, and the particle dispersion system composed of the two phases reaches a good dynamic equilibrium, and the reaction is more stable.

[0054] Preferably, the interior of the pipe reactor 210 needs to be polished. A pipe with a large surface roughness will disrupt the flow state of the fluid, causing friction with the pipe wall, thereby generating a large friction resistance. Preferably, the surface roughness of the pipe reactor 210 is less than Ra0.2μm.

[0055] Preferably, the two-phase fluids converge in the collision zone ③, which is characterized in that the inlet directions of the continuous phase ① and the continuous phase ② form a certain angle θ1, which is 0 to 180 degrees, preferably 0 to 120 degrees, and more preferably 30 to 90 degrees.

[0056] Preferably, in the above reaction process, the diffusion zone ④ is a columnar structure, the number of which is 0-10, preferably 2-5. Preferably, the columnar structure is an S-shaped channel or a zigzag channel.

[0057] Preferably, in the above reaction process, the length L of the diffusion zone is a multiple of the cross-sectional diameter of the diffusion zone, L / R, which is 3-50, preferably 3-10 times.

[0058] Preferably, the throttle structure in the above reaction process is quasi-circular, preferably circular.

[0059] Preferably, the throttle cross-sectional angle θ2 is 60-150 degrees, preferably 90-120 degrees.

[0060] Preferably, in the above reaction process, the cross-sectional area S2 of the throttle port is 0.3-1.5 times, preferably 0.5-0.8 times, the sum of the cross-sectional areas S1 of all diffusion zones.

[0061] Preferably, the mixers in the above reaction process can be used in parallel, or three stages can be used in parallel, or two stages can be used in series or in parallel.

[0062] Preferably, the reaction process further comprises a pipe reactor 210, which cannot be used alone. The length of the pipe reactor is adjustable, and the inner surface roughness of the pipe reactor is less than Ra0.2 μm.

[0063] Preferably, the pipe reactor described in the above reaction process is internally provided with a mixing unit. The mixing unit is a hollow blade structure. The outer edge of the mixing unit blade closely fits the pipe wall and extends inward in a hyperbolic shape along the radial direction of the pipe. After reaching the center, it continues to extend to the pipe wall to form a mixing unit blade. This guides the fluid to exchange positions with the boundary layer of the pipe wall in the central region, thereby eliminating dead zones. Preferably, the reaction unit described in the above reaction process also includes a separator, which is a pressurized separation device.

[0064] Preferably, the reaction unit in the above reaction process further includes a separator 300, and the separator 300 is a dynamic separation device or a static separation device.

[0065] Preferably, the separator in the above reaction process is an automatic separation device. The two-phase fluid outlet pipeline is provided with a regulating valve, the opening of which is adjusted at any time according to the interface meter installed in the separator to maintain a stable position of the two-phase interface.

[0066] Preferably, the reaction process further comprises an energy integration unit, which utilizes the sensible heat generated after the pressure release of the flash evaporator to achieve thermal energy conversion of the fluid and separation of the two-phase fluid. The energy integration unit comprises a heat exchanger, characterized in that the vapor formed after the flash evaporation exchanges energy with the liquid phase, thereby utilizing the energy generated by the reaction.

[0067] Preferably, in the above reaction process, the continuous phase after separation in the flash evaporator passes through a distributor 320 , which is connected to two outlets. The distributor can automatically adjust the flow rate, one of which is connected to the mixing kettle 110 and the other is connected to the waste acid tank 360 .

[0068] Preferably, in the above reaction process, raw materials and a continuous phase are added to a mixing kettle 110, and the continuous phase is fed to the continuous phase inlet of the mixer via a delivery pump 111. The dispersed phase in the raw material tank is also fed to the dispersed phase inlet of the mixer via a pump. The continuous phase is circulated back from the distributor to the mixing kettle, and the ratio of the two phases is adjusted to a volume ratio of 1:0.5 to 1:30, preferably 1:2 to 1:12.

[0069] The application of the above-mentioned reaction process of the present invention in chemical reaction processes preferably includes nitration, etherification, oxidation, esterification, amination, chlorination, and diazotization. The process can be applied to adiabatic reactions as well as heat transfer reactions.

[0070] In the above preferred technical solution, the reaction process of the present invention has a more excellent effect.

[0071] The present invention is further illustrated below through the accompanying drawings and detailed description of specific embodiments, but is not intended to limit the present invention and is merely provided for illustrative purposes.

[0072] Example 1 Tubular reaction process for producing 1000 tons of mononitro-p-dichlorobenzene per year

[0073] A concentrated sulfuric acid solution is used as a continuous phase fluid, and the sulfuric acid is used as a circulating stream;

[0074] Using p-dichlorobenzene as dispersed phase fluid;

[0075] The device adopts negative pressure adiabatic reaction type and uses a single-stage mixer for the reaction.

[0076] refer to Figure 2 , 63.4 kg / h of fresh 98 wt% concentrated sulfuric acid is added to the mixing kettle 110 through pipeline 1110, and the concentrated sulfuric acid in the circulating stream returned from the downstream system is fully mixed in the mixing kettle. The concentrated sulfuric acid circulation flow rate in the return stream is 766.8 kg / h, and the concentration of the concentrated sulfuric acid in the circulating stream is 78.7 wt%. After mixing, it is continuously fed into the mixer 200 as a continuous phase through the delivery pump 111. 53.3 kg / h of concentrated nitric acid, the reaction raw material, is added at the inlet of the delivery pump 111, with a concentrated nitric acid concentration of 98 wt%. The delivery pump 111 is a high-pressure turbine mixing pump. The continuous phase concentrated sulfuric acid and the raw material concentrated nitric acid are vigorously mixed under the centrifugal force of the high-speed rotation in the pump. After sufficient mixing and dissolution, the mixed acid enters the mixer 200 at high pressure. The temperature of the pipeline 1112 entering the mixing kettle is set to 75 ° C and the flow rate is set to 883.7 kg / h. The organic phase raw material p-dichlorobenzene of the reaction enters the other end of the mixer 200 as a dispersed phase through the pipeline 1000, the flow rate is set to 115.3 kg / h, and the temperature of the reaction raw material entering the mixer pipeline 1000 is set to 50°C.

[0077] Figure 6 Schematic diagram of the mixer structure. The mixed acid enters the mixer 200 at high pressure from the continuous phase inlet in the schematic diagram, and the pressure is set to 18 bar; the organic phase raw material enters the mixer 200 from the dispersed phase inlet in the schematic diagram, and the pressure is set to 8 bar.

[0078] After entering the mixer, the two-phase mixed stream is fed into tubular reactor 210, with a residence time of 130 seconds. The length of the two-phase stream in the tubular reactor can be adjusted arbitrarily, and the residence time is controllable, allowing for precise regulation and optimization of the reaction process. Mixer 200 and tubular reactor 210 form the entire reactor, and the reaction is adiabatic. The heat released during the reaction causes the temperature of the stream in the tubular reactor to rise sharply, reaching a maximum of 100°C after mixer 200.

[0079] After a full reaction, the two-phase flow system enters separator 300. Kinetic energy from the high-pressure pump is not fully released in the reactor and enters separator 300 under pressure. Separator 300 is equipped with an interface meter. This meter, combined with the two-phase outlet regulating valve, ensures a stable interface between the two phases. The two-phase flow system consists of immiscible phases, which automatically separate due to their density differences. The denser circulating sulfuric acid phase flows out from the bottom 3001 at a rate of 843.4 kg / h. The less dense mononitro product organic phase flows out from the top 3002 at a rate of 155.6 kg / h.

[0080] The huge amount of heat released after the adiabatic reaction causes the system temperature to rise, and the pressure of the fluid drops instantly after passing through the regulating valve. The circulating sulfuric acid phase enters the flash evaporator 310 after decompression from the bottom 3001, and the sensible heat in the sulfuric acid fluid is released. The water generated after the nitration reaction reduces the concentration of concentrated sulfuric acid, and the sensible heat in the fluid can evaporate part of the water into the gas phase. This waste heat can increase the concentration of the circulating sulfuric acid and be recycled. The energy released by the reaction is fully utilized and recycled, and the cost of the product is effectively reduced. The sulfuric acid concentration increases after negative pressure flash evaporation. About 15kg / h of water is evaporated, and the flash evaporation pressure is set to 25kPa (absolute pressure).

[0081] The flash vapor is transferred from pipeline 3101 to the organic phase of the mononitro product separated from separator 300 through heat exchanger 330 for further waste heat recovery. It condenses into water and enters wastewater tank 350. After exchanging heat with the organic product phase, it enters heat exchanger 340 via pipeline 3300. The sulfuric acid flowing from flash evaporator 310 continues to transfer heat with the heated organic product phase through heat exchanger 340. The sulfuric acid from distributor 320 exchanges heat with the organic product phase, then enters mixing kettle 110 through pipeline 3403 for recycling. The organic product phase enters crude product tank 370, with an organic phase flow rate of 154.2 kg / h.

[0082] In this embodiment, the mixer is a "contraction-diffusion-contraction" collision structure, and the concentrated sulfuric acid mixed acid enters the Figure 6 The continuous phase inlet ① is shown. Through flow restriction, the continuous phase fluid completes its first “contraction”. The dispersed phase enters Figure 6 The continuous phase is at the inlet of dispersed phase. Figure 6 The energy is released for the first time at the collision zone ③. Due to the kinetic energy difference between the continuous phase and the dispersed phase, the small droplets formed after the continuous phase is atomized collide with the dispersed phase at high speed, the interface between the two phases is highly intertwined and mutually soluble, and the two phases are efficiently mixed here. Figure 6The diffusion zone ④ is shown here. The channel area is enlarged and the two-phase fluid undergoes a "diffusion" process here. The diffusion zone can be set to different shapes and lengths, and the reaction continues. After passing through the diffusion zone, the fluid reaches the throttle. Control Figure 6 Figure 5 shows the throat cross-sectional area at throttle port ⑤, where the two-phase mixture undergoes its second "contraction." The kinetic energy of the two-phase flow continues to transform at throttle port ⑤, where the fluids continue to shear and mix, completing the second half of the reaction. By controlling the cross-sectional area of ​​the throttle port, the flow velocity of the two-phase mixture can reach subsonic levels. The entire process of the continuous and dispersed phases' immiscible flows within the mixer completes a "contraction-diffusion-contraction" process, achieving energy transfer and thus material transformation.

[0083] In this embodiment, Figure 7 θ1 shown is 90 degrees; Figure 6 The number of diffusion zone ④ is 2, and the diffusion zone channel is linear; Figure 8 θ2 shown is 120 degrees; Figure 10 S2 is shown as 0.5 times the sum of S1.

[0084] In this example, concentrated sulfuric acid, seven times the mass of the raw material, was used as the circulating stream to ensure that the temperature rise in the logistics system remained within a safe range, i.e., it did not exceed TD24, the temperature at which the maximum reaction rate of the runaway reaction, TMRad, is reached after 24 hours, as specified in the "Reaction Risk Assessment." In the reaction risk assessment, TMRad—the maximum temperature the system can reach after a runaway reaction—is 92.3°C, and TD24 is 123.7°C. In this example, using an adiabatic reaction, the reaction system temperature reached a maximum of 105°C, which is lower than TD24—the initiation temperature after a 24-hour adiabatic induction period, which is 125.7°C.

[0085] The heat released by the reactor reaction is absorbed by the material system within the reactor, resulting in an adiabatic temperature rise of 30°C. The maximum temperature of the reaction system is approximately 25°C below the triggering temperature of a runaway reaction. The maximum temperature the system can reach is far from the TD24 triggering temperature, ensuring safe production of the process unit from the very nature of the reaction.

[0086] The specific production data of the process product of Example 1 is shown in Table 1-1.

[0087] Table 1-1

[0088]

[0089]

[0090] Comparative Example 1: Production capacity is 1000 tons of mononitro-p-dichlorobenzene per year

[0091] The difference between Comparative Example 1 and Example 1 is that Comparative Example 1 adopts a microchannel nitration reaction process. The production data of the microchannel nitration reaction process are shown in Table 1-2.

[0092] Table 1-2

[0093]

[0094] Comparing the production data table with Example 1, it can be seen that the dinitro content in the product is greatly increased by adopting the microchannel nitration reaction process, and the raw material content in the product is large, which does not meet the requirements for using the product as raw material to produce downstream products.

[0095] In Example 1, the process of the present invention can achieve full utilization of reaction energy, utilize negative pressure flash evaporation to evaporate the water generated in nitration production, and realize the recycling of continuous phase sulfuric acid.

[0096] This production process can save the construction cost and high operating costs of the factory's waste acid concentration unit. It fully utilizes the potential of flash steam to achieve low-cost and high-efficiency operation of the entire production process.

[0097] Furthermore, it is well known that the selectivity of the nitrobenzene reaction decreases significantly with increasing conversion. In actual production, the dinitrobenzene content can rise sharply to 1.2%, which is unacceptable for high-quality mononitrobenzene products. The reaction process of the present invention compensates for the "underreaction" of the main reaction and "overreaction" of the side reaction in the mixer reactor.

[0098] Example 2 Production capacity is 9,000 tons per year of m-chloro-3-nitrotrifluorotoluene

[0099] refer to Figure 3 In Example 2, concentrated sulfuric acid solution is used as the continuous phase fluid, and the sulfuric acid is used as a circulating stream; m-chlorotrifluorotoluene is used as the dispersed phase fluid;

[0100] The device adopts adiabatic reaction type and adopts series and parallel two-stage mixers.

[0101] The specific process of Example 2 is similar to that of Example 1. The difference is that:

[0102] 660 kg / h of fresh 98 wt% concentrated sulfuric acid is added to the mixing kettle 110 and fully mixed with the circulating stream of concentrated sulfuric acid returned from the downstream system distributor 320 in the kettle. The concentrated sulfuric acid circulation rate in the return stream is 3700 kg / h, and the concentration of the circulating stream concentrated sulfuric acid is 85.9 wt%. After mixing, it is continuously fed into the mixer 200 as a continuous phase through the delivery pump 111. 403 kg / h of 98 wt% concentrated nitric acid, the reaction raw material, is added to the inlet of the delivery pump 111. After sufficient mixing and dissolution, the mixed acid enters the mixer 200 under high pressure, with the pressure set to 12 bar. This embodiment uses a two-stage mixer to control the flow rate. The flow rate of pipeline 1113 entering the primary mixer 200 is set to 1500 kg / h, and the flow rate of pipeline 1112 entering the secondary mixer 201 is set to 3300 kg / h. The temperature of pipelines 1112 and 1113 is set to 60°C.

[0103] The organic phase raw material m-chloro-trifluorotoluene was fed as the dispersed phase through pipe 1000 into the other end of the mixer. The pressure was set to 4 bar, the flow rate was set to 1083 kg / h, and the temperature of the raw material entering the mixer pipe 1000 was set to 50°C.

[0104] To prevent overreaction, which can lead to a sharp increase in isomeric impurities, a secondary mixer is provided. The two-phase mixture first reacts in the primary mixer 200. After reaction and mixing, it continues to react with the secondary dispersed phase in the secondary mixer 201. This compensates for the "underreaction" of the primary reaction and the "overreaction" of the side reaction in the mixer reactor.

[0105] After entering the mixer, the two-phase mixture flows into tubular reactor 210, with a residence time of 150 seconds. The reaction is adiabatic, and the heat released during the reaction causes the temperature of the stream in the tubular reactor to rise sharply, reaching a maximum of 100°C after the secondary mixer 201.

[0106] In this embodiment, Figure 7 θ1 shown is 60 degrees; Figure 6 The number of diffusion zone ④ is 5, and the diffusion zone channel is S-shaped; Figure 8 θ2 is shown as 90 degrees; Figure 10 S2 is shown as 0.6 times the sum of S1.

[0107] In this example, concentrated sulfuric acid, 4.7 times the mass of the raw material, was used as the circulating stream to ensure that the temperature rise in the logistics system remained within a safe range, i.e., did not exceed TD24, the temperature at which the maximum reaction rate of the runaway reaction, TMRad, is reached after 24 hours, as specified in the "Reaction Risk Assessment." In the reaction risk assessment, TMRad—the maximum temperature the system can reach after a runaway reaction—is 125°C, and TD24 is 146°C. In this example, using an adiabatic reaction, the reaction system temperature reached a maximum of 100°C, well below TD24—the initiation temperature of 146°C after a 24-hour adiabatic induction period.

[0108] The heat released by the reactor reaction is absorbed by the material in the reactor itself. The adiabatic temperature rise of the system is 40°C. The highest temperature that the system can reach is far away from TD24, ensuring the safe production of the process from the essence of the reaction.

[0109] After the reaction, the two-phase system is pressurized and enters separator 300. The two-phase system is an insoluble phase, and due to the difference in density, it automatically separates. The denser circulating sulfuric acid phase flows out from the bottom 3001, and the less dense mononitro product organic phase flows out from the top 3002. The organic phase flow rate is 1349 kg / h.

[0110] Sulfuric acid flowing from the flash evaporator enters distributor 320 via pipeline 3001. One stream flows through 3201 into waste acid tank 360 at a rate of 770 kg / h. Another stream flows through pipeline 3202, passes through heat exchanger 340, and continues to transfer heat with heated organic product phase 3300. After being heated, it enters mixing kettle 110 via pipeline 3403 for recycling. The rate of sulfuric acid is 3700 kg / h.

[0111] The concentration of sulfuric acid solution increases after negative pressure flash evaporation. About 20kg / h of water is evaporated, and the flash pressure is set at 15kPa (absolute pressure).

[0112] After heat exchange, the organic product phase enters the crude product tank 370, and the flow rate of the organic phase is 1337 kg / h.

[0113] The specific production data of the process product of Example 2 is shown in Table 2-1.

[0114] Table 2-1

[0115]

[0116] Comparative Example 2:

[0117] The production capacity is 9,000 tons of p-chloro-3-nitrotrifluorotoluene per year. The difference between Comparative Example 2 and Example 2 is that the two-stage mixer of the present invention is not used, but a conventional tubular reaction process is adopted. The production data are shown in Table 2-2.

[0118] Table 2-2

[0119]

[0120] Comparing Example 2 and Comparative Example 2, the production data table shows that the conventional tubular reaction process results in a significant increase in the dinitro content in the product, which is unsatisfactory for using the product as a raw material for producing downstream products. In order to avoid excessive dinitro content in the product, this embodiment employs a two-stage mixer in series and parallel to complete production, thus avoiding the situation in which excessive dinitro content would result from overreaction in a first-stage reaction. The entire production process is controllable.

[0121] In this process, concentrated sulfuric acid in the logistics system is used as the circulating continuous phase, and the mass ratio of the dispersed phase to the continuous phase is 1:4.7. Under this condition, sufficient volume heat capacity is guaranteed, the adiabatic temperature rise of the reaction is controlled within 40 degrees, and the temperature is far away from the triggering temperature of the reaction out of control, ensuring the inherent safety of production.

[0122] The process production device can realize the full utilization of the secondary energy used in the reaction, and use the negative pressure flash evaporation to evaporate the water generated in the nitration production to realize the recycling of the continuous phase sulfuric acid.

[0123] The process in Example 2 fully utilizes reaction energy, using negative pressure flash evaporation to evaporate some of the water generated during nitration. This process reduces the high operating costs of the spent acid concentration unit. By leveraging waste heat technology to fully exploit the steam potential, the entire production process achieves low-cost and high-efficiency operation.

[0124] Due to the increase in production capacity, the use of microchannels has been limited. The initial investment of using microchannel technology is huge, which is completely disadvantageous in terms of project investment compared with the process provided by the present invention. In today's fierce market competition, minimizing costs while ensuring product quality is the trend of enterprise development.

[0125] Table 2-3 below compares the investment in various process equipment and production energy consumption costs for the construction of a nitration synthesis production unit with an annual output of 9,000 tons of m-chloro-3-nitrotrifluorotoluene.

[0126] Table 2-3 Unit: Ten thousand yuan

[0127]

[0128] The invention has simple process flow, advanced equipment, small investment, and fully utilizes the energy balance in the reaction process, reducing the consumption of energy and labor costs, thereby reducing the overall production cost.

[0129] Example 3 Production capacity is 5000 tons of p-nitroanisole per year

[0130] refer to Figure 4In Example 3, methanol solution is used as the continuous phase fluid, and the methanol is used as a circulating stream;

[0131] 3,4-Dichloronitrobenzene is used as the dispersed phase fluid;

[0132] The process adopts heat transfer reaction type and the reaction adopts single-stage mixer.

[0133] The specific process of Example 3 is similar to that of Example 1. The difference is that:

[0134] Sodium methoxide (1073.9 kg / h) was added to mixing kettle 110 via pipeline 1110. A recycle stream of methanol (1780.9 kg / h) was returned from the downstream system via heat exchanger 340 and thoroughly mixed in the mixing kettle. After mixing, the mixture was continuously fed to mixer 200 via delivery pump 111 as the continuous phase. The mixture entered mixer 200 under high pressure, set at 8 bar. The temperature of pipeline 1112 entering the mixing kettle was set at 60°C.

[0135] The organic phase raw material 3,4-dichloronitrobenzene is fed as the dispersed phase into the other end of the mixer via pipe 1000. The pressure is set at 2 bar, the flow rate is set at 775.5 kg / h, and the temperature of the raw material entering the mixer pipe 1000 is set at 65°C.

[0136] The two-phase mixed flow enters the mixer and then enters the tubular reactor 210. The residence time of the flow is set to 215 seconds. The reaction adopts heat transfer reaction, and the reaction temperature is controlled at 75-80°C during the reaction.

[0137] In this embodiment, Figure 7 θ1 shown is 30 degrees; Figure 6 The number of diffusion zone ④ is 4, and the diffusion zone channel is a broken line shape; Figure 8 θ2 shown is 100 degrees; Figure 10 S2 is shown as 0.8 times the sum of S1.

[0138] After the reaction, the two-phase flow system enters the separator 300 under pressure. The two-phase flow system is a slurry system, and due to the difference in density, it can automatically separate. The dense salt slurry heavy phase flows out from the bottom 3001, and the light etherified mixed product phase flows out from the top 3002 at a flow rate of 2338.5 kg / h and enters the flash evaporator 310.

[0139] After the etherified mixed product phase was flashed under negative pressure, about 322 kg / h of methanol and water were evaporated. The flash pressure was set to 10 kPa (absolute pressure), and the methanol content of the flash gas phase was 97.7%.

[0140] After the vaporized phase exchanges heat with the salt slurry heavy phase through pipeline 3101, the salt slurry enters the salt slurry tank 360. The solid content of the salt slurry is 32.6%. After condensation, the methanol vapor returns to the mixing kettle 110 through 3403 and is recycled as the continuous phase.

[0141] In Example 3, the advantages of the present invention are fully utilized for industrial devices that produce solid slurries in reaction synthesis systems, particularly in microchannel synthesis processes, which are severely limited. By fully utilizing reaction energy and using negative pressure flash evaporation to evaporate methanol as the continuous phase fluid, the process reduces the operating costs of methanol refining, achieving low-cost and high-efficiency operation of the entire production device.

[0142] Example 4 Production capacity is 1000 tons of p-nitrobenzoic acid per year

[0143] refer to Figure 5 , using acetic acid as a continuous phase fluid, and the acetic acid is used as a circulating stream;

[0144] p-Nitrotoluene is used as the dispersed phase fluid;

[0145] The device adopts adiabatic reaction type and adopts series and parallel two-stage mixers.

[0146] The specific process of Example 4 is similar to that of Example 2. The difference is that:

[0147] 2449.0 kg / h of fresh glacial acetic acid is added to the mixing kettle 110, and the circulating acetic acid returned from the downstream flash evaporator 310 is fully mixed in the kettle. The returned logistics is in the gas phase and is inserted into the bottom of the mixing kettle to provide heat for the logistics system. The return logistics flow rate is 522.5 kg / h, the concentration of the circulating logistics acetic acid is 83.5 wt%, and the temperature is 111.3 ° C. After mixing, it is continuously input into the mixer 200 as a continuous phase through the delivery pump 111. 386.3 kg / h of 30 wt% dilute nitric acid, a reaction raw material, is added to the inlet end of the delivery pump 111. After sufficient mixing and dissolution, the mixed acid enters the mixer under high pressure, and the pressure is set to 12 bar.

[0148] This embodiment uses a two-stage mixer to control the flow rate. The flow rate of pipeline 1113 entering the first-stage mixer 200 is set to 971.5 kg / h, the flow rate of pipeline 1112 entering the second-stage mixer 201 is set to 2000 kg / h, and the temperature of pipelines 1112 and 1113 is set to 120°C.

[0149] The organic phase raw material p-nitrotoluene of the reaction enters the other end of the mixer 200 as the dispersed phase through the pipeline 1000. The flow rate is set to 120.6 kg / h, the pressure is set to 3 bar, and the temperature of the reaction raw material entering the mixer pipeline 1000 is set to 100°C.

[0150] After entering the mixer, the two-phase mixture is fed into tubular reactor 210, with a residence time of 310 seconds. The reaction is adiabatic. The heat released during the reaction causes the temperature of the stream in the tubular reactor to rise sharply, reaching a maximum of 180°C after the secondary mixer 201.

[0151] In this embodiment, Figure 7 The mixing angle θ1 is 30 degrees; Figure 6 The number of diffusion zone ④ is 3, and the diffusion zone channel is linear; Figure 9 、 Figure 10 The L / R shown is 15; Figure 8 θ2 is shown as 90 degrees; Figure 10 S2 is shown as 0.5 times the sum of S1.

[0152] In this example, a sufficient amount of acetic acid is used as a circulating stream to ensure that the temperature rise of the logistics system remains within a safe range. The heat released by the reactor reaction is absorbed by the materials within the reactor, resulting in an adiabatic temperature rise of 70°C. The maximum temperature that the system can reach is well above TD24, ensuring safe production of the process unit based on the nature of the reaction.

[0153] After a full reaction, the mixed-phase flow system enters the flash evaporator 310 under pressure. The vapor phase flowing out of the flash evaporator flows back to the mixing kettle 110 via 3403, with a vapor phase temperature of 111.3°C. The latent heat of vapor liquefaction is utilized in the mixing kettle 110. The NO generated by the reaction is also returned to the mixed liquid, effectively reducing overall costs. The liquid phase flows through 3300 to the crude product tank 370.

[0154] The process production apparatus of Example 4 utilizes flash evaporation technology to directly recycle vaporized gas, thereby utilizing reaction energy. In particular, nitric oxide (NO) gas is returned to the system, reducing waste gas treatment costs, effectively controlling raw material quota consumption, and lowering product costs. This achieves low-cost and high-efficiency operation of the entire production apparatus.

[0155] In this process, concentrated sulfuric acid in the logistics system is used as the circulating continuous phase, and the mass ratio of the dispersed phase to the continuous phase is 1:4.5. Under this condition, sufficient volume heat capacity is guaranteed, the adiabatic temperature rise of the reaction is controlled within 60 degrees, and the temperature is far away from the triggering temperature of the reaction out of control, ensuring the inherent safety of production.

[0156] In order to reflect the influence of the mixer structure parameters on the reaction, different data θ1, θ2 and L / R are compared, and the corresponding production data are shown in Table 4-1.

[0157] Table 4-1

[0158]

[0159] It can be seen from Table 4-1 that different structural parameters in the mixer have an impact on the reaction results.

[0160] The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the technical concept of the present invention, various simple variations of the technical solution of the present invention may be made, including combining the various technical features in any other appropriate manner. These simple variations and combinations should also be regarded as disclosed in the present invention and fall within the scope of protection of the present invention.

Claims

1. An intrinsically safe, efficient tubular reaction process, characterized in that: include: a) a raw material preparation unit, comprising a mixing kettle (110) for batching, a continuous phase feed pipeline (1110), a raw material pipeline (1111), a delivery pump (111), a delivery pipeline (1112) and a raw material tank (100); b) a reaction unit, comprising a mixer (200), a pipeline reactor (210), a separator (300), a separator light phase pipeline (3002), and a separator heavy phase pipeline (3001); c) an energy integration unit, comprising a flash evaporator (310), a flash evaporator gas phase pipeline (3101), a distributor (320), a heat exchanger (330), a heat exchanger (340), a heat exchanger outlet pipeline 3300 and a batching kettle (110) return pipeline (3403); d) A post-processing unit, comprising a waste water tank (350), a waste acid tank (360) and a crude product tank (370).

2. The process according to claim 1, characterized in that Raw materials and a continuous phase are added to a mixing kettle (110). The raw materials are added from the top of the mixing kettle (110) or from the inlet of a delivery pump (111). The delivery pump (111) is preferably a high-pressure turbine mixing pump.

3. The process according to claim 1 or 2, characterized in that The mixer (200) is a "contraction-diffusion-contraction" type collision structure, and its structure mainly consists of a continuous phase inlet ①, a dispersed phase inlet ②, a collision zone ③, a diffusion zone ④ and a throttling port ⑤.

4. The process according to any one of claims 1 to 3, characterized in that The continuous phase enters from the continuous phase inlet ① at high pressure, with a pressure of 0-25 bar, preferably 3-18 bar; the raw material enters from the dispersed phase inlet ② at low pressure, with a pressure of 0-10 bar, preferably 1-6 bar; And / or the continuous phase and the dispersed phase two-phase fluids merge in the collision zone ③, and the continuous phase and the continuous phase inlet direction form a certain angle θ1, and the angle θ1 is 0 to 180 degrees, preferably 0 to 120 degrees, and more preferably 30 to 90 degrees; And / or the diffusion zone ④ is a columnar structure, the number of the columnar structures is 0-10, preferably 2-5; preferably, the columnar structure is a straight channel, an S-shaped channel or a folded channel; And / or the longitudinal length L of the diffusion zone ④ is a multiple of the cross-sectional diameter R of the diffusion zone, L / R is 3-50 times, preferably 10-30 times. and / or the throttle structure is quasi-circular, preferably circular; and / or the cross-sectional angle θ2 of the throttle port ⑤ is 60-150 degrees, preferably 90-120 degrees; And / or the cross-sectional area S2 of the throttling port ⑤ is 0.3-1.5 times, preferably 0.5-0.8 times, the sum of the cross-sectional areas S1 of all diffusion zones.

5. The process according to any one of claims 1 to 4, characterized in that The mixer (200) can be used alone, or in two-stage cascade, or in three-stage cascade. And / or the mixer (200) can be used in parallel, or three stages can be used in parallel, or two stages can be used in series and in parallel.

6. According to the process according to any one of claims 1 to 5, the pipeline reactor (210) cannot be used alone, characterized in that: The pipeline reactor (210) can be used in combination with a conventional reactor, such as the pipeline reactor (210) can be used in combination with a microchannel reactor; preferably, the length of the pipeline reactor can be adjusted to change the residence time of the logistics system, and the inner surface roughness of the pipeline reactor is lower than Ra0.2μm.

7. The process according to any one of claims 1 to 6, characterized in that A mixing unit is arranged inside the pipe reactor (210), and the mixing unit is a hollow blade structure. The outer edge of the mixing unit blade is tightly matched with the pipe wall and extends inward in a hyperbolic shape along the radial direction of the pipe. After reaching the center, it continues to extend to the pipe wall to form a hyperbolic mixing unit blade; preferably, the hyperbolic mixing unit is arranged in sequence in a 45-degree direction, and the number of blades is three or more.

8. The process according to any one of claims 1 to 7, characterized in that The separator (300) is a pressurized or normal pressure separation device, and the separator (300) is a dynamic separation device or a static separation device; preferably, the separator (300) is an automatic, static separation device, and two-phase interface regulating valves are provided on the two-phase fluid outlet pipelines of the continuous phase and the dispersed phase.

9. The process according to any one of claims 1 to 8, characterized in that After separation in the flash evaporator (310), the continuous phase passes through a distributor (320), which is connected to two outlets. The two outlets of the distributor (320) can automatically distribute the flow, one of which returns to the mixing kettle (110), and the other is connected to the waste acid tank (360).

10. The process according to any one of claims 1 to 9, characterized in that A raw material and a continuous phase are added to a mixing kettle (110), and the raw material and the continuous phase enter the continuous phase inlet ① of a mixer (200) through a delivery pump (111). The dispersed phase in a raw material tank (100) enters the dispersed phase inlet ② of a mixer (200) through a pump (1000). The continuous phase is circulated back to the mixing kettle (110) from a distributor (320), and the ratio of the two phases is adjusted. The volume ratio of the dispersed phase to the continuous phase is 1:0.5 to 1:30, preferably 1:2 to 1:

12.

11. Use of the process according to any one of claims 1 to 10 in an adiabatic reaction or a heat transfer reaction.

12. The use according to claim 11, characterized in that: The applications include nitration, etherification, oxidation, esterification, amination, chlorination and diazotization.

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