CN1112913A - Method of technical cumene hydroperoxide acidic cleavage to phenol acetone and alpha-methylstyrene - Google Patents

Abstract

Technical grade cumenyl hydroperoxide with wide composition range in out-of-band The cleavage is carried out in a tubular circular reactor with a shell, using and dicumyl Oxides are at the same high pressure in a columnar continuous reactor with a circulation ratio of at least for 26. The pyrolysis process is controlled by two calorimeters, i.e. the control of the absolute ΔT value, which is the difference in temperature reduction between the two calorimeters.

Description

This invention relates to the synthesis of phenol, acetone and alpha-methylstyrene (AMS) using cumene (cumene).

The cumene process consists of two stages: the first stage is the oxidation of cumene to cumene hydroperoxide (CHP) with oxygen in the air, and the second stage is the acidic catalytic cracking (decomposition) of CHP to phenol and acetone. During CHP cracking, certain unwanted products called "phenol tars" are formed along with phenol, acetone and AMS. The amount of phenol tar mainly depends on the CHP cracking method used, in the best current process, the amount is 50-60kg/t phenol, in the traditional process it is 120-180kg/t phenol.

To date, many skilled artisans have endeavored to increase the selectivity or increase the yield of the process. However, output or nominal output is an important indicator as well as selectivity when measuring the overall production capacity of a phenol-acetone production facility. So far it has not been possible to simultaneously satisfy the above two requirements.

Recent technological advances are based on the CHP cleavage method described in USP No. 5,254,751. The industrial grade CHP cracking carried out according to this method is shown in Fig. 1 . The method is accomplished in two steps.

The first step is the cracking of CHP catalyzed by sulfuric acid. This step is carried out in three shelled tubular reactors A, B, and C connected in series, and the heat of reaction is taken away by cooling water added to the gap between the reactor tubes. . The cleavage product is circulated in the reaction unit at a ratio of 10-25/1 of the circulating amount to the CHP added amount. The most convenient way to add sulfuric acid catalyst to the reactor is to add it to the circulating cleavage product through line 14, pump 2 sends the circulating cleavage product and catalyst sulfuric acid to the mixer 16, and fresh technical grade CHP passes through the line 17 is also fed to mixer 16, from which this mixture is sent to reactor A. In order to increase the selectivity of the method, another amount of water and acetone is added to the circulating cracking stream, depending on the amount of CHP added, the added acetone is calculated according to the following algorithm (1):

C _ac =G _CHP) ×0.17[CHP]+ 40/(G _CHP [CHP])

Where: Gac-Acetone content, t/hr

G _CHP - the amount of industrial grade CHP, t/hr

[CHP] - the weight concentration of CHP in technical grade CHP.

When building each algorithm listed in this specification, its quantity is measured in metric tons (ton or t). However, any other unit of weight may be used as long as the consistent use of the same unit of quantity is chosen for each component.

Those skilled in the art know that technical grade CHP is impure cumene hydroperoxide with varying amounts of impurities such as DMBA, cumene, AP and the like, which are products of cumene oxidation , and the oxidation of cumene is the first step in the preparation of phenol from cumene.

The process of circulating acetone is very important for the selectivity of the method, which is determined by the temperature range in each of the three reactors (A, B, C in the first section of the first step in the method) , were strictly controlled at 50-62°C, 62-57°C and 57-50°C under atmospheric pressure, and the conversion of CHP in each reactor was 30-60%, 25-50% and 30-10%, respectively. Reactors A, B, C are vented to atmosphere through appropriately designed vents 13 .

Because the cracking reaction of CHP is exothermic, it is necessary to use a reactor with a total specific heat exchange area of not less than ^30-35m2 per metric ton of 100% CHP. Because of the potential danger of CHP cracking reaction, it is actually safer to design a reactor with an exchange area of not less than 45-60m ² per metric ton of 100% CHP feed ratio.

The CHP decomposition process in the first step is monitored in the system through a specially designed calorimeter 3 (small reactor) installed on the outlet pipeline of the last reactor. The temperature between the inlet and outlet materials of the calorimeter The difference (δT ₁ ), is a quantitative measure of the amount of undecomposed CHP remaining in the cleavage product. Typical ? _T? values are 4-16°C, which correspond to 0.6-2.3 wt% of free CHP present in the last reactor of the first step of the process. At the same time, during the cracking process of CHP, the impurity dimethylbenzyl alcohol (DMBA) present in the raw materials of industrial-grade CHP will partially react with CHP at the same time to generate the intermediate dicumenyl peroxide (DCP).

Within the above-mentioned range of circulation ratio and CHP conversion rate, the CHP cleavage time is from 30 seconds to 3 minutes, preferably 45 seconds to 2 minutes.

The cleavage of DCP in the second step is done continuously in two unidirectional flow reactors 4 and 7, and ammonium hydroxide aqueous solution is added to the first reactor 4 through line 15 to convert a part of sulfuric acid into NH ₄ HSO ₄ . Therefore, the cracking of DCP is completed in the reactor 7 at a temperature of 90-110°C by using a binary catalyst (H ₂ SO ₄ +NH ₄ HSO ₄ ) with a certain component control ratio, and the product is heated by heat exchanger 6 to complete.

In the reactors A, B and C of the first stage, the cracking of CHP is carried out at atmospheric pressure and vented to the atmosphere through appropriately designed vent holes 13 . In reactor 4, the cracking of DCP is accomplished under atmospheric pressure, and in reactor 7, it is accomplished at not less than 2 atm. A pressure reducing valve 22 is installed between the reactor 7 and the evaporator 8 .

A part of the acetone in the cleavage product is evaporated in the evaporator 8 under atmospheric pressure or about subatmospheric pressure, preferably at an absolute pressure of 500-600 mmHg, and the remaining cleavage product leaves the evaporator from the bottom and is sent to a pipeline 20 for further processing into Products phenol and acetone. The evaporated acetone is cooled in cooler 9 and concentrated into a liquid, and the liquid acetone is delivered to vessel 10 before being pumped by pump 11 into reactors A, B and C of the first stage.

The method described above gave the highest level of method selectivity prior to the present invention, but there are several areas where significant improvements could be made:

1) A reactor with a large capacity is required to provide a large amount of heat exchange (more than 50m ² /t 100%CHP), which will lead to high equipment prices and high capital investment in construction and equipment compensation capabilities.

2) The intermediate vessel (reactor 4) operates at atmospheric pressure and requires a pump 5 to transfer the lysate to the DCP converter 7 of the 2nd step of the process.

3) The reduction of CHP cracking temperature in the first section will lead to a large amount of undecomposed CHP (from 0.5-2%) in the reactor 4. If the ratio of H ₂ SO ₄ : NH ₄ HSO ₄ is wrong, it will There is danger.

4) It is difficult to accurately measure a small amount of ammonia water and H ₂ SO ₄ in order to maintain the correct ratio. In the case of inaccurate metering and incomplete neutralization of H ₂ SO ₄ , AMS dimers will be generated and phenol complexes will be formed. , in the case of excessive neutralization of H ₂ SO ₄ , the conversion of DCP will be incomplete. In both cases the yield of the desired product is reduced and the "phenol tar" production is naturally increased.

5) H ₂ SO ₄ and NH ₄ HSO ₄ present in the cleavage products, will cause the acidic catalytic properties of these catalysts to increase 4-7 times, and their concentration in vessel 8 will increase simultaneously during the evaporation of acetone, in which Side reactions continue in the vessel and reduce the selectivity of the process due to the occurrence of by-products. The aforementioned low dose of NH ₄ OH exacerbates the situation and leads to a loss of 10-20 kg of raw material cumene for the production of 1 ton of phenol, and the production of AMS will also drop to 60% of the theoretical amount.

6) To operate the DCP cracking reactor 7 with a very weak acidic catalyst (H ₂ SO ₄ +NH ₄ HSO ₄ ), a relatively high reactor volume is required, that is, 0.8m ³ /t is required for each ton of the reactor, Compared with the reactor volume of 0.4-0.5m ³ /t per ton of feed in the method of the present invention, it is very unfavorable.

7) The acetone added to the CHP cracking equipment through the above algorithm (1) will lead to higher energy requirements for its concentration, and equipment with improved heat exchange capacity is required, which also leads to a large amount of money spent on equipment for adjustment and compensation .

On the contrary, the present invention increases the production capacity of the process equipment, reduces capital investment by reducing equipment volume when building new equipment, and simplifies the process, but still maintains high process selectivity.

The prior art method is shown in FIG. 1 . The method of the present invention is shown in FIG. 2 .

In the method of the present invention, the cracking steps of CHP and DCP are completed under the same pressure in the reactors connected in series. Wherein the pressure in the reactor is about 2-10 atm. Preferably the pressure is about 3-5 atm. CHP is added to the recycle loop of the cracking reaction product containing the sulfuric acid catalyst. As in the prior art method, the sulfuric acid catalyst is usually added to the recycle loop through line 14 and technical grade CHP is added through line 17 and mixed in mixer 16 with the circulating CHP lysate. In the prior art method, reactors A, B, and C are open to the atmosphere, but in the method of the present invention, line 21 is connected to the top of reactors A, B, C, and 7, and keeps all reactors in common. at the same higher pressure. Because all four reactors A, B, C and 7 are at the same pressure, and reactor 4 is deleted in the process of the present invention, the ammonia solution added through line 15 in the prior art process of Fig. 1 is shown , is added to the line entering the evaporator 8 after the pressure reducing valve 22 in this method. Although it is most convenient to add ammonia as shown in Figure 2 before the nearest evaporator 8, and indeed there will be no ammonia present in the reactor 7, as long as there is no ammonia in the reactor 7, the addition position is after the reactor 7 , it doesn't matter whether it is added before or after the pressure reducing valve 22.

In the present invention, the investment in pumps and reactor 4 is not required since both the first and second steps of the process are designed to operate at the same higher pressure. The first step reaction of the present invention is carried out at a higher temperature than in the prior art method, and this higher temperature requires the use of higher pressure in order to lower the boiling point of the lysate. The high pressure of the first step of the method of the present invention is the same as that of the second step, so no pump is needed.

With the cracking of CHP, the heat released is taken away by the cooling water in the reactors A, B and C of the first stage of the process in the first stage of the process. The weight ratio of reactant circulation rate to CHP feed rate is not less than about 26:1, but not more than about 40:1, preferably from about 26:1 to about 35:1.

According to the method of the present invention, join the acetone in the 1st workshop section reactor A of the 1st step of this method, and its ratio to the industrial CHP quantity that adds for cracking is calculated by following algorithm (II):

C _ac =G _CHP) ×0.125[CHP]+ 35/(G _CHP [CHP])

Where: Gac- is the amount of acetone added for cracking, t/hr

G _CHP - Amount of technical grade CHP added for cracking, t/hr

[CHP] - the weight concentration of CHP in technical grade CHP.

This formula is applicable to the most commonly available industrial CHP, such as industrial grade CHP with a CHP weight concentration of about 74%-92%.

Broader algorithms applicable to a wider range have been discovered and provide the basis for a method to control the use of a wider range of technical-grade CHP feedstocks.

Algorithm (Ⅲ) is as follows:

G _ac ＝0.58G _CHP × ((1.315[CHP])/(50+0.25[CHP]) - (2[CHP])/152 - ([DMBA])/136 - ([cumene])/120 - ([AP])/120 )

Where: Gac- is the amount of acetone added for cracking, t/hr

G _CHP - Amount of technical grade CHP added for cracking, t/hr

[CHP] - CHP weight concentration in technical grade CHP.

[DMBA] - DMBA weight concentration in technical grade CHP.

[cumene] - cumene weight concentration in technical grade CHP.

[AP] - weight concentration of acetophenone (AP) in technical grade CHP.

The above algorithm extends the applicability of the invention to CHP feedstocks with CHP concentrations as low as 40% by weight and as high as about 98% by weight, preferably about 50-90%, more preferably about 60-85%, but can also be used to use any Method for CHP concentration.

In the present invention, high temperature and higher conversion of CHP are employed in the first step of the CHP decomposition process. Reactor A, the conversion range of B and C are controlled at 55-78%, 60-94% and 90-98% respectively, and the total conversion rate of CHP is close to complete in the present invention. Calorimeter 3 (miniature reactor) is installed at the outlet of reactor C, which gives an indication of the overall conversion of CHP by the above-mentioned δT ₁ value. Typical values for _δT1 are from about 0.4-2.1°C at pressures from about 3 to 4 atmospheres, typical temperature ranges for reactors A, B and C are 57-82°C, 65-82°C and 57-70°C, respectively .

The concentration of CHP at the reactor outlet remains on 0.1-0.45wt%, preferably 0.2-0.4wt%, and the cracking time of CHP is from 17 seconds to 28 seconds. In the method of the present invention, the total specific surface area of heat exchange is 17- 25m ² /t 100% CHP reactor.

The lysate from reactor C passes through heater 6 into reactor 7 where DCP and DMBA are converted to the desired products.

A sufficient quantity of water to provide 98% conversion of DCP and DMBA in reactor 7 was added via line 20 to stationary mixer 19 on the reaction C product outlet line. The content of water in the cracked product is controlled at no more than 3wt%, preferably 1.3-2.0wt%.

Controlling the degree of decomposition of DCP and DMBA after the water and cracking product mixing position is achieved by monitoring the temperature difference ( _δT2 ) of the second small reactor 12 (calorimeter) in the system, and the small reactors 12 are connected in parallel On the pipeline connecting the reactors A, B and C of the first section of the first step of the method to the reactor 7. The temperature in the mini-reactor 12 is controlled at both the inlet and the outlet.

The configuration of calorimeter 12 is not critical as long as the residence time of the cracked products in calorimeter 12 is sufficient for the overall conversion of DCP and DMBC to phenol, acetone and AMS. Generally this can be achieved by low velocity flow of the cracked product slip stream.

The temperature control of this method is accomplished by temperature difference (δT), which is the difference between δT ₁ of the first calorimeter and δT ₂ of the second calorimeter, and the absolute value of δT is maintained at 0.2- within 3°C.

In order to exclude chemical losses during the evaporation of the acetone as the product enters the evaporator 8, a neutralizing agent base is added via line 18 in an amount which must completely neutralize the sulfuric acid as a neutral salt (sulphate). Na ₂ CO ₃ , NH ₄ OH, and NaOH can also be used as neutralizing agents, and ammonia water with a concentration of 1-10wt% is better.

The method of the present invention has the following advantages:

1) The scope of application of the method is wider, it can adopt industrial-grade CHP (40-98% by weight) with a wider range of CHP concentration, and has high processing efficiency.

2) The production capacity of the equipment can be increased by 2-2.5 times without a large amount of capital investment. While the output is high, the selectivity still maintains a maximum value (the yield of AMS is 78-80% of the theoretical amount).

3) 50-60% reduction in investment in new crackers to be built due to:

a) The CHP cracking reaction occurs in a reactor with a low specific heat exchange area (17-25m ² /t CHP), while the specific heat exchange area of a traditional reactor is 40-60m ² /t CHP.

b) The DCP cracking reaction occurs in one reactor instead of two reactors, because the two reactors can be operated at the same pressure, so the two reactors can be compacted into one device.

4) Reduced energy consumption in operation due to reduced amount of recycled acetone and elimination of pump 5 from the process.

5) The addition of ammonia to the DCP cracking step is omitted, and only water is added, which simplifies the control of the acidic catalytic performance of H ₂ SO ₄ .

6) The process control of δT=δT ₂ -δT ₁ simplifies the CHP cracking process, and automatically helps to maintain the selectivity of the method in the CHP cracking stage and DCP cracking stage.

7) Completely neutralizing _{the H2SO4} prior to the acetone vaporizer helped rule out unwanted side reactions.

In addition, the leakage of low-boiling acetone in the existing method is eliminated by connecting the reactors A, B and C of the first workshop section and the DCP cracking reactor in the method realizing the new process.

The invention is illustrated in detail by the following examples 2-14, without being limited by these examples, and compared with comparative example 1 which represents the prior art.

Comparative Example 1

Using the prior art method shown in Figure 1, technical grade CHP with the following wt% composition was continuously added to a reactor unit consisting of three shelled tubular reactors with a total reaction capacity of 10.08 m ³ :

Cumene hydroperoxide 82.500wt%

Cumene 12.721wt%

Dimethyl benzyl alcohol 4.325wt%

Acetophenone 0.453wt%

26 tons of industrial grade CHP per hour are added to CHP reactors A, B and C, according to the calculation of algorithm (1), 5492kg/h of acetone is added, 16kg/h of sulfuric acid is added to the above reactor, and The cleavage products are recycled therein. For the cracking reaction, the total heat transfer surface of the reactor is 1254 m ² , which corresponds to a specific heat transfer value of 58 m ² /t CHP per ton of 100% CHP.

The residence time in the CHP cleavage reactor was 74 seconds. The temperature on the outlet pipeline of each reactor is respectively 58 ℃, 55 ℃ and 50 ℃, and the CHP conversion rate at each reactor outlet is respectively 38%, 73% and 85%, and the temperature reduction value in the small reactor 3 ( δT ₁ ) is 5.6°C, CHP is mixed with the reacted lysate at the inlet of the cleavage reactor at a ratio of 1:16 (the circulation ratio is 16), and 10kg of sulfuric acid is added to the circulation loop.

The cracking of DCP is accomplished in two unidirectional flow reactors working continuously at low temperature. The first reactor temperature was 58°C and the second was 93°C. 33.8 kg/h of ammonia solution is fed to the first reactor at a concentration of 5 wt%, so the degree of neutralization of sulfuric acid for the conversion of sulfuric acid to ammonium bisulfate will be 50%.

The cracking of CHP in three shelled pipeline reactors and the cracking of DCP in two unidirectional flow reactors were carried out at different pressures, atmospheric pressure in the CHP cracking reactor and the first reactor of DCP cracking, DCP The pressure of the second reactor for cleavage was 3-5 atm.

The residence times in the DCP cleavage reactors were 420 seconds in reactor 4 and 2030 seconds in reactor 7, respectively.

The acetone is transported from the container 10 to the cracking device through the pump 11, the above-mentioned acetone is evaporated from the reaction lysate of DCP under reduced pressure in the evaporator 8, and concentrated into a liquid in the cooler 9.

The overall reaction results that occur in the reactor are:

Phenol- 13174.5kg/h (99.2% yield)

Acetone- 8084.5kg/h (98.9% yield)

α-Methylstyrene - 54.5kg/t phenol (AMS yield is 73.4% of theoretical amount considering added DMBA)

Phenol tar production - 59.2kg/t phenol

Cumene consumption- 1333kg/t phenol

In this embodiment and subsequent examples, the consumption value of cumene is only for comparing the effect of illustrating the method of the present invention, and the present invention only relates to a part of the whole phenol method. When other cumene recovery operations not part of the present invention are included in the cumene consumption in the overall process, the net cumene consumption is even lower, typically 1307-1310 kg/t.

Example 2

This example of the invention demonstrates a 2.3-fold higher yield than Example 1. The method of the present invention in this embodiment and the following embodiments is completed as described in FIG. 2 . All reactors of this example were operated at the same pressure, and the operating pressure of the whole system was 4 kg/cm ² .

The technical grade CHP of 60 tons per hour is added in the CHP cracking reactor, and its composition is identical with embodiment 1. The cracking of industrial-grade CHP is carried out under the following conditions, the circulation ratio is 26/1, the temperature at the inlet of each reactor is 68 ° C, 67 ° C and 60 ° C, and the cumulative CHP conversion rate in each reactor is 62 %, 94% and 98%.

According to the calculation of the formula (II), the acetone concentrated from the container 10 is added to the cracking equipment (the first section) at a speed of 6890kg/h:

C _ac =G _CHP) ×0.125[CHP]+ 35/(G _CHP [CHP])

Among them, Gac is the amount of acetone, and G _CHP is the amount of industrial-grade CHP, both of which are expressed in t/hr, where 1 ton (t) is equal to 1000 kilograms (kg). [CHP] is the concentration of CHP in industrial-grade CHP, expressed in wt%, and the temperature drop value in the small reactor 3 of the cracking equipment is δT ₁ =0.4°C.

By entering the pipeline of DCP cracking reactor 7, water is added at a speed of 482kg/h, so that the concentration of water in the cracking product at the reactor outlet is 1.63wt%, and the cracking time is 485 seconds.

At the point where the cleavage product and water are mixed, a small reactor 12 (calorimeter) is installed where the temperature drop ( _δT₂ ) between the outlet and the inlet is measured. Process control is actually achieved by absolute temperature reduction,

δT = δT ₂ -δT ₁ = 0.65°C

Aqueous ammonia solution at 10% wt concentration and a rate of 145 kg/h was added via line 18 to the product inlet line into the acetone evaporator 8 in order to exclude chemical losses of the desired product in the acetone evaporator 8.

the result is:

Phenol- 30431.2kg/h (99.5% yield)

Acetone- 18769.7kg/h (98.5% yield)

α-Methylstyrene - 58.5kg/t phenol (AMS yield is 79.9% of theoretical amount considering added DMBA)

Phenol tar production - 54.2kg/t phenol

Cumene consumption- 1328kg/t phenol

Example 3

This example illustrates the process of the invention carried out in reduced capacity reactors A, B and C which are reduced heat transfer surface and low specific surface heat exchangers. This process is completed according to the route described in embodiment 2, but the equipment used in embodiment 3 is:

The volume of CHP cracking reactor system is smaller than embodiment 1 2.8 times,

The surface of this system heat exchanger is smaller than embodiment 1 3.5 times,

The specific surface area of the DCP-cracked heat exchanger was 3 times smaller than the similar value described in Example 1.

26t/hr of industrial grade CHP with the same composition as Example 1 is added to the CHP cracking reactor, and 9.3kg/hr of sulfuric acid is added to the CHP cracking reactor (A, B, C in the first section) In the circulation loop of , the cracking of industrial grade CHP is carried out under the following conditions. The temperatures at the outlets of reactors A, B and C are 79°C, 75°C and 69°C respectively, and the conversion rates of CHP are 77%, 96% and 98% respectively. %.

4312 kg/h of circulating acetone calculated by formula II in Example 2 is fed from vessel 10 to the CHP cleavage unit.

The residence time of the reaction product in DCP cleavage reactor 7 was 638 seconds.

the result is:

Phenol- 13185.3kg/h (99.4% yield)

Acetone- 8075.4kg/h (yield 98.8%)

α-Methylstyrene - 58.02 kg/t phenol (AMS yield 79.9% of theory considering added DMBA)

Phenol tar output- 55.03kg/t phenol

Cumene consumption- 1329kg/t phenol

In the small reactor 3 of the cracking plant, the temperature drop is δT ₁ =0.34°C.

Absolute temperature difference δT = δT ₂ -δT ₁ = 0.66°C

Concentration is that the aqueous ammonia solution of 10%wt is added in the product inlet line of acetone evaporator 8 with the speed of 63kg/h, so that get rid of the chemical loss of desired product in acetone evaporator, so sulfuric acid is converted into sulfuric acid in acetone evaporator 8 The degree of ammonium conversion was 100%.

Example 4

This example illustrates the practice of the invention in reactors A, B and C where the heat exchange specific surface area is determined to be 25 m ² /t 100% CHP.

Composition is determined by the technical grade CHP of 26t/hr of embodiment 1, is added in the CHP cracking reactor, and circulation ratio is 26/1. The sulfuric acid of 9.3kg/h is added by CHP cracking reactor A, B and C In the composed circulation loop, the cracking of CHP is carried out at the outlet temperatures of reactors A, B and C respectively at 67°C, 66°C and 61°C, and the conversion rates of CHP are 62%, 87% and 94%, respectively.

Acetone circulated at a rate of 4312 kg/h, calculated by the formula II proposed in Example 2, was fed from vessel 10 to the cleavage unit of CHP.

The residence time of the reaction product in the DCP cracking reactor 7 is 640 seconds.

The result is the following quantity:

Phenol- 13188.5kg/h (99.3% yield)

Acetone- 8078.9kg/h (yield 98.8%)

α-Methylstyrene - 58.5 kg/t phenol (AMS yield 79.9% of theory considering added DMBA)

Phenol tar production - 53.9kg/t phenol

Cumene consumption- 1328kg/t phenol

In the mini-reactor 3 of the cracking plant, the temperature drop _δT1 = 1.26°C, the mini-reactor 12 installed on the outlet of the circulation loop after water is added to the reaction product, the temperature drop in the mini-reactor 12 is 1.98°C.

The absolute temperature difference (δT = δT ₂ -δT ₁ ) is 0.72°C

Concentration is the aqueous ammonia solution of 10%wt, passes through pipeline 18 with the speed of 63kg/h, adds in the product inlet pipeline that leads to acetone evaporator 8, so that get rid of the chemical loss of desired product in acetone evaporator 8, so in The degree of conversion of sulfuric acid to ammonium sulfate in the acetone evaporator is 100%.

Example 5

This example illustrates the relatively low rate of feed of technical grade CHP per cracker and the excellent selectivity of the process of the present invention at high recycle ratios.

Composition is joined in the CHP cracking reactor by the technical grade CHP of 22t/h specified in Example 1, and the circulation ratio is 35, and in the circulation loop that is made up of CHP cracking reactor A, B and C, add the sulfuric acid of 7.2kg/h . The cracking of industrial grade CHP was carried out at the outlet temperatures of reactors A, B and C at 72°C, 78°C and 67°C, respectively, and the conversion rates of CHP were 65%, 92% and 97%, respectively.

The 4192kg/hr circulating acetone calculated by formula II in Example 2 is added to the cracking equipment (I section) of CHP from the container 10.

The residence time of the reaction product in DCP cleavage reactor 7 was 865 seconds.

In the small reactor 3 of the cracking plant, the temperature drop value δT ₁ =0.82°C

In the small reactor 12 installed after the means for supplying water at the outlet of the reaction product circulation loop, the temperature drop was 1.68°C.

The absolute temperature difference (δT=δT ₂ -δT ₁ ) is 0.86°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet pipeline entering the acetone evaporator 8 at a rate of 48kg/h, so as to avoid the chemical loss of the desired product in the acetone evaporator, so that sulfuric acid becomes the conversion degree of ammonium sulfate In the acetone evaporator 8 it is 100%.

The following quantities of product were obtained as a result:

Phenol- 11156.4kg/h (99.4% yield)

Acetone-6815.7kg/h (98.5% yield)

α-Methylstyrene - 58.1 kg/t phenol (AMS yield 79.9% of theory considering added DMBA)

Phenol tar production - 55.14kg/t phenol

Cumene consumption- 1328kg/t phenol

Example 6

This example illustrates the high selectivity of the process of the present invention with the addition of technical grade CHP to each cleavage unit at a moderate rate.

Composition is made up of the 35t/h technical grade CHP specified by embodiment 1, is added in the CHP cracking reactor, and the recirculation ratio is 26, adds the sulfuric acid of 11.80kg/h in the circulation loop that is made up of CHP cracking reactor. The cracking of industrial grade CHP was carried out at the outlet temperatures of reactors A, B and C at 74 °C, 71 °C and 65 °C, respectively, and the conversion rates of CHP were 70%, 93% and 97%, respectively.

According to the calculation formula II among the embodiment 2, the circulating acetone of 4821kg/h is added to the CHP cracking equipment (the first section) from the container 10.

The residence time of the reaction product in DCP cleavage reactor 7 was 487 seconds.

In the mini-reactor 3 of the cleavage plant, the temperature drop is δT ₁ =0.64°C, and in the mini-reactor 12 installed after the device for supplying water to the reaction products at the outlet of the circulation loop, the temperature drop is 1.92°C.

The absolute temperature difference (δT = δT ₂ -δT ₁ ) is 1.72°C

Concentration is that the ammoniacal solution of 10% is with the speed of 80kg/h, is added to the product inlet pipeline that enters acetone evaporator 8 by pipeline 18, so that prevent the chemical loss of desired product in acetone evaporator, so that sulfuric acid becomes the generation of ammonium sulfate The degree of conversion was 100% in the acetone evaporator.

The results obtained are as follows:

Phenol- 17751kg/h (99.4% yield)

Acetone- 10906kg/h (98.5% yield)

α-Methylstyrene - 58.6 kg/t phenol (AMS yield 79.9% of theory considering added DMBA)

Phenol tar production - 54.14kg/t phenol

Cumene consumption- 1328kg/t phenol

Example 7

This example demonstrates the highly selective reproducibility of the process at moderate feed rates of technical grade CHP per cracker. In this example, the concentration of technical grade CHP is actually higher than that in Comparative Example 1 and Examples 2-6.

35 t/h of technical grade CHP with the following composition was added to the CHP cracking reactor.

CHP 91.5% (wt)

Cumene 2.0% (wt)

DMBA 5.5% (wt)

Acetophenone 1.0% (wt)

The circulation ratio is 26, and the sulfuric acid of 9.3kg/h is added in the reaction loop that is made up of CHP cracking reactor A, B and C. The cracking of industrial grade CHP was completed when the reactor outlet temperatures were 71°C, 67°C and 61°C, respectively, and the conversion rates of CHP were 75%, 94% and 98%, respectively.

The circulation acetone of 4444kg/h calculated by the formula II in the embodiment 2 is added to the CHP cracking equipment from the container 10.

The residence time of the reaction products in the DCP cleavage reactor 7 was 640 seconds.

In the small reactor 3 of the cracking plant, the temperature drop value δT ₁ =0.49°C, and the temperature drop value in the small reactor 12 installed at the outlet of the circulation loop after adding water to the reaction product line is 1.38°C.

The absolute temperature difference (δT=δT ₂ -δT ₁ ) is 0.89°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet line leading to the acetone evaporator 8 at a rate of 63 kg/h to prevent the chemical loss of the desired product in the acetone evaporator 8, so that the conversion of sulfuric acid into ammonium sulfate The degree is 100% in the acetone evaporator.

and get the following result:

Phenol- 14610.2kg/h (99.4% yield)

Acetone- 8972.3kg/h (98.5% yield)

α-Methylstyrene - 56.6 kg/t phenol (AMS yield 79.9% of theory considering added DMBA)

Phenol tar production - 62.51kg/t phenol

Cumene consumption- 1335kg/t phenol

The above examples illustrate the extremely high selectivity of the process at various levels of feed rate for technical grade CHP under various reaction conditions with excellent reproducibility.

Example 8

The present embodiment and the following examples (9-14) illustrate that the output of Example 1 is increased by 2.3 times through the preferred embodiment of the present invention. It uses the formula III to calculate the amount of acetone added, and illustrates that in industrial grade CHP there is an extremely Using this embodiment is also effective over a broad concentration range of 40-98%. Compared with Example 2, this method can be completed without changing the configuration of the equipment. All reactors in the process of the present invention are operated at the same pressure, so all reactors are connected to each other through the top, and the total pressure of the system is 4 kg/cm ² .

From the concentrated acetone of container 10, add in the cracking equipment with the speed 6890kg/h calculated by following formula (Ⅲ):

G _ac ＝0.58G _CHP × ((1.315[CHP])/(50+0.25[CHP]) - (2[CHP])/152 - ([DMBA])/136 - ([cumene])/120 - ([AP])/120 )

Where: Gac- is the amount of acetone added for cracking, t/h

G _CHP - the amount of technical grade CHP added for cracking, t/h

[CHP] - the weight concentration of CHP in industrial grade CHP

[DMBA] - the weight concentration of DMBA in industrial grade CHP

[Cumene] - the weight concentration of cumene in technical grade CHP

[AP] - weight concentration of acetophenone in technical grade CHP

The technical grade CHP of 60t/h adds in the CHP cracking reactor A, B and C, and its composition is identical with embodiment 1. 21.3kg/h of sulfuric acid is also added to the reactor, and the cracking of industrial grade CHP is carried out at the reactor inlet temperature of 68°C, 67°C and 60°C at a circulation ratio of 26/1, which is equivalent to a CHP conversion rate of 59 %, 94%, 98%.

In the small reactor 3 of the cracking plant, the temperature drop is δT ₁ =0.4°C.

Water is added to the product feed line entering DCP cracking reactor 7 through line 20 and fixed mixer 19 at a rate of 455 kg/h, so that the concentration of water in the cracked product at the reactor outlet is 1.58% ( wt), the lysis time was 485 seconds.

A second small reactor 12 (calorimeter) is installed at the position after mixing the pyrolysis products and water, and the temperature drop δT ₂ at its outlet and inlet is measured, the process control is essentially done by the absolute temperature drop:

δT = δT ₂ -δT ₁ = 0.62°C

Aqueous ammonia solution with a concentration of 10% is passed through line 18 at a linear velocity of 141 kg/h and added to the product inlet line entering the acetone evaporator 8, so as to eliminate the chemical loss of the desired product in the acetone evaporator 8, causing the acid to become sulfuric acid The degree of conversion of ammonium in the acetone evaporator 8 is 100%.

and get the following result:

Phenol- 30434.1kg/h (99.5% yield)

Acetone- 18774.8kg/h (98.5% yield)

α-Methylstyrene - 58.4kg/t phenol (AMS yield is 78.9% of theory considering the added DMBA)

Phenol tar production - 54.0kg/t phenol

Cumene consumption- 1328kg/t phenol

Example 9

This example illustrates the implementation of the process according to the invention when the capacity of the reactor is reduced, the heat exchanger surface of the reactor is reduced, and the specific surface of the heat exchanger is therefore lower. This process is carried out according to the route described in embodiment 8, but the equipment used in this embodiment is as follows:

The capacity of CHP cracking reactor system is smaller than embodiment 1 2.8 times,

The heat exchanger surface of this system is 3.5 times smaller than that of embodiment 1,

The specific surface area of the DCP cracking heat exchanger is 3 times smaller than the corresponding value described in Example 1,

Form the technical grade CHP of 26t/h as specified in embodiment 1, be added in the A, B and C reactor of CHP cracking, to the circulation system (the 1st working section) that is made up of CHP cracking reactor A, B and C ) with 9.3kg/h of sulfuric acid. The cracking of industrial-grade CHP is completed when the outlet temperatures of reactors A, B and C are 73°C, 70°C, and 65°C, respectively, corresponding to CHP conversion rates of 66%, 90%, and 96%.

The 4312kg/h acetone calculated by the formula III among the embodiment 8 is added to the CHP cracking equipment (the 1st workshop section) from the container 10.

The residence time of the reaction product in the DCP cleavage reactor 7 was 630 seconds.

and get the following result:

Phenol- 13183.2kg/h (99.4% yield)

Acetone- 8072.5kg/h (yield 98.8%)

α-Methylstyrene - 58.02 kg/t phenol (AMS yield is 78.4% of theory considering added DMBA)

Phenol tar production - 55.43kg/t phenol

Cumene consumption- 1329kg/t phenol

In the small reactor 3 of the cracking plant, the temperature drop δT ₁ is 0.95°C.

Absolute temperature difference δT = δT ₂ -δT ₁ = 0.58°C

Concentration is that the aqueous ammonia solution of 10% is added in the acetone evaporator 8 product inlet pipelines by pipeline 18 with the speed of 63kg/h, so that get rid of the chemical loss of desired product in acetone evaporator 8, so that sulfuric acid changes into ammonium sulfate The degree of conversion was 100% in the acetone evaporator 8.

Example 10

This example illustrates the implementation of the process of the invention in a reactor based on 100% CHP with a heat exchange specific surface area of 25 m ² /t. Form the technical grade CHP of 26t/h specified by embodiment 1, add in the CHP cracking reactor, circulation ratio is 26, add 9.3kg/h in the circulation loop that is made up of CHP cracking reactor A, B and C sulfuric acid. The cracking of industrial grade CHP was carried out at the outlet temperatures of A, B and C reactors at 67°C, 66°C and 60°C, respectively, corresponding to CHP conversions of 62%, 82% and 90%, respectively.

The acetone of 4313kg/h calculated by the formula III among the embodiment 8 is added to the CHP cracking equipment from the container 10.

The residence time of the reaction products in the DCP cleavage reactor 7 was 640 seconds.

The resulting data is as follows:

Phenol- 13188.4kg/h (99.3% yield)

Acetone- 8082.9kg/h (98.8% yield)

α-Methylstyrene - 50.1 kg/t phenol (AMS yield 79.8% of theory considering added DMBA)

Phenol tar production - 52.9kg/t phenol

Cumene consumption- 1327kg/t phenol

In the mini-reactor 3 of the cleavage plant, the temperature drop δT ₁ is 2.17°C, and in the mini-reactor 12 installed at the outlet of the circulation loop after supplying water to the reaction products, the temperature drop is δT ₂ =3.11°C.

The absolute temperature difference (δT=δT ₂ -δT ₁ ) is 0.94°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet line of the acetone evaporator 8 through a line 18 at a rate of 63 kg/h, so as to prevent the chemical loss of the desired product in the acetone evaporator, so that the conversion of sulfuric acid into ammonium sulfate The degree is 100% in the acetone evaporator.

Example 11

This example illustrates the high selectivity of the process of the present invention under the conditions of lower feeding rate and higher circulation ratio of industrial-grade CHP for each cracking device.

The technical grade CHP of 22t/h that composition is stipulated by embodiment 1 is added in the CHP cracking reactor A, B and C. The cycle ratio is 35. 8.3 kg/h of sulfuric acid are added to the circulation loop consisting of CHP cleavage reactors A, B and C. The cracking of industrial grade CHP was carried out at the outlet temperatures of reactors A, B and C at 72°C, 78°C and 67°C, respectively, and the conversion rates of CHP were 68%, 94% and 98% respectively.

The speed calculated by embodiment 8 formula III is the acetone of 4191.4kg/h, adds to CHP cracking equipment (the 1st workshop section) from container 10.

The residence time of the reaction products in DCP cleavage reactor 7 was 737 seconds.

The temperature drop in the small reactor 3 of the cracking plant is δT ₁ =0.39°C, and the temperature drop in the small reactor 12 installed at the outlet of the circulation loop after supplying water to the reaction product is 0.99°C.

The absolute temperature difference (δT=δT ₂ -δT ₁ ) is 0.6°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet line of the acetone evaporator 8 through a line 18 at a rate of 55.1 kg/h, so as to eliminate the chemical loss of the desired product in the acetone evaporator, so that sulfuric acid becomes ammonium sulfate The degree of conversion in the acetone evaporator 8 was 100%.

and get the following result:

Phenol- 11159.1kg/h (yield 79.4%)

Acetone- 6821.8kg/h (98.5% yield)

α-Methylstyrene - 57.8kg/t phenol (AMS yield is 78% of theory considering added DMBA)

Phenol tar production - 55.6kg/t phenol

Cumene consumption- 1328kg/t phenol

Example 12

This example illustrates the selectivity of the process of the present invention when each cracking device is fed with technical grade CHP at a moderate rate.

The 35t/hr industrial grade CHP whose composition is specified in Example 1 is added to CHP cracking reactors A, B and C. The cycle ratio was 26. Add 13.1 kg/h of sulfuric acid to the circulation loop consisting of CHP cleavage reactors A, B and C. Cracking of technical grade CHP was carried out at reactor A, B and C outlet temperatures of 73°C, 71°C and 66°C, respectively, corresponding to CHP conversions of 55%, 82% and 91%.

According to the calculation formula III among the embodiment 8, the circulating acetone of 6410kg/h is added to the cracking equipment (the 1st workshop section) from the container 10.

The residence time of the reaction product in DCP cleavage reactor 7 was 467 seconds.

In the mini-reactor 3, the temperature drop was _δT1 = 2.09°C, and in the mini-reactor 12 installed on the outlet of the circulation loop after water supply to the reaction product, the temperature drop was 3.04°C.

The absolute temperature difference (δT=δT ₂ -δT ₁ ) is 0.95°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet line of the acetone evaporator 8 at a rate of 88.7kg/h, so as to eliminate the chemical loss of the desired product in the acetone evaporator 8, so that the conversion of sulfuric acid into ammonium sulfate The degree is 100% in the acetone evaporator.

and get the following result:

Phenol- 17752.4kg/h (99.4% yield)

Acetone- 10904.3kg/h (99.1% yield)

α-Methylstyrene - 58.9 kg/t phenol (AMS yield 79.67% of theory considering added DMBA)

Phenol tar production - 53.84kg/t phenol

Cumene consumption- 1328kg/t phenol

The above-mentioned embodiment 12 illustrates the practicality of the present invention (i.e. 35t/h and 26t/h for comparison) under the situation that a relatively high level of material is passed through the CHP cracking equipment

The following examples illustrate the applicability of the present invention using technical grade CHP feedstocks of varying CHP concentrations. Example 13 shows that the CHP concentration in the raw material is very high, and Example 14 shows that the CHP in the raw material is poor.

Example 13

This example of technical grade CHP is used to illustrate the high selectivity of the process of the invention at high concentrations of CHP.

26t/hr technical grade CHP of following composition is added in the CHP cracking reactor A, B and C:

CHP 91.5% (wt)

Cumene 2.0% (wt)

DMBA 5.5% (wt)

Acetophenone 1.0% (wt)

The recycle ratio is 26, and 9.3 kg/h of sulfuric acid is added to the recycle loop consisting of CHP cleavage reactors A, B and C. Cracking of industrial grade CHP was carried out at reactor A, B and C outlet temperatures of 72.5°C, 68°C and 61°C, corresponding to CHP conversions of 78%, 96% and 98%.

According to the calculation formula III among the embodiment 8, the circulating acetone of 4444kg/h is added to the CHP cracking equipment (the 1st workshop section) from the container 10.

The residence time of the reaction product in reactor 7 was 642 seconds.

In the small reactor 3 of the cracking plant, the temperature drop value δT ₁ was 0.36°C, and in the small reactor 12 installed at the outlet of the circulation loop after supplying water to the reaction product, the temperature drop value was 1.46°C.

The absolute temperature difference (δT = δT ₂ -δT ₁ ) is 1.1°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet line of the acetone evaporator 8 at a rate of 63 kg/h, so as to eliminate the chemical loss of the desired product in the acetone evaporator 8, so that the conversion of sulfuric acid into ammonium sulfate The degree is 100% in the acetone evaporator.

The result is the following data:

Phenol- 14607.4kg/h (99.3% yield)

Acetone- 8975kg/h (99.9% yield)

α-Methylstyrene - 64.8kg/t phenol (AMS yield is 77.33% of theory considering added DMBA)

Phenol tar production - 69.36kg/t phenol

Cumene consumption- 1342kg/t phenol

Example 14

This example illustrates the high selectivity of the process of the invention at low concentrations of CHP.

26t/hr technical grade CHP of following composition is added in the CHP cracking reactor A, B and C:

CHP 67% (wt)

Cumene 28.6% (wt)

DMBA 4% (wt)

Acetophenone 0.4% (wt)

The recycle ratio is 26, and 11.8 kg/h of sulfuric acid is added to the recycle loop consisting of CHP cleavage reactors A, B and C. The cracking of CHP was carried out at the outlet temperatures of reactors A, B and C at 69°C, 63°C and 62°C, respectively, and the conversions of CHP were 68%, 96% and 98% respectively.

According to the speed calculated by formula III in embodiment 8, the circulating acetone of 2376kg/h is added to the CHP cracking reaction equipment (the first section) from the container 10.

The residence time of the reaction product in DCP cleavage reactor 7 was 516 seconds.

The temperature drop value δT ₁ was 1.58°C in the small reactor 3 and 2.04°C in the small reactor 12 installed on the outlet of the circulation loop after water supply to the reaction product.

The absolute temperature difference (δT=δT ₂ -δT ₁ ) is 0.46°C

Aqueous ammonia solution with a concentration of 10% is added to the product inlet line of the acetone evaporator 8 through the line 18 at a rate of 79 kg/h, so as to eliminate the chemical loss of the desired product in the acetone evaporator 8, so that the sulfuric acid becomes ammonium sulfate The degree of conversion of is 100% in the acetone evaporator 8.

The following results are obtained:

Phenol- 14402kg/h (99.2% yield)

Acetone- 8833kg/h (yield 98.9%)

α-Methylstyrene - 67kg/t phenol (AMS yield 79.9% of theory considering DMBA added)

Phenol tar production - 79.9kg/t phenol

Cumene consumption- 1331kg/t phenol

Below is the comparison of the method of the present invention and USP No.5,254,751 method operation.

Compared with the method of USP No.5,254,751, the cracking method of the present invention is carried out faster and more "reliable". Three reactors A, B and C in the method of the present invention are smaller, and are operated at higher temperatures (and higher pressures to lower the boiling point)

temperature °C

A B C pressure

The present invention 68-79 78-65 69-60 3-4atm

US5,254,751 50-62 62-57 57-50 1atm

The conversion of CHP through the three reactors is different from USP No. 5,254,751. In the process of the present invention, a higher percentage of the CHP raw material is consumed by the reaction and leaves the first reactor, practically no unreacted CHP comes out of the reactor C, and the _δT of the calorimeter is controlled at a very low value ( <1°C).

%CHP Conversion % wt% CHP

A B C Calorimeter 3

The present invention 75 90 98 0.2

US5,254,751 45 75 88 1.0

In the process of the present invention, a higher recycle ratio (26:1-40:1) is used in reactors A, B, and C, so that the CHP feedstock gets additional dilution, so safe operation can be maintained at a faster cracking rate . The CHP concentration added to reactor A is 2.5-3 wt%, while in USP No. 5,254,751 method, this value is 4.5-5 wt%. CHP residence time is 17-28 seconds in the 1st step of the present invention, and is 50-60 seconds in the method of US 5,254,751.

For the first step reaction, the concentrations of sulfuric acid, water and acetone were substantially the same in both processes and the concentration of sulfuric acid catalyst was 300 ppm in both cases.

In the method of the present invention, the outlets of the first step and the second step are connected together, so the two steps are operated under the same elevated pressure, which saves the investment in pumps and other equipment.

The above examples, experiments and expected results are presented to illustrate and illustrate the invention and are not intended to limit the invention to those parameters specifically disclosed. Therefore, many improvements described above can be made by carefully reading this specification, but as defined by the following claims, these should be counted as the scope and content of the present invention.

The data excerpts for the examples are in the table below:

Claims (12)

1. An improved method for the production of phenol and acetone from cumene, the method comprising acid catalytic cracking technical grade cumenyl hydroperoxide and cracking dicumenyl peroxide, wherein the improvement comprises: in the first reaction Cracking technical grade cumenyl hydroperoxide in a reactor to obtain cumenyl hydroperoxide lysate containing dicumyl peroxide and the same Under pressure, dicumyl peroxide is cracked, so no pumps and make-up equipment are needed to transfer the cumyl hydroperoxide cleavage from the cumyl hydroperoxide cleavage reactor to the dicumyl peroxide cleavage reactor . 2. The process according to claim 1, wherein the pressure in the cumenyl hydroperoxide cleavage reactor and the dicumenyl peroxide cleavage reactor is about 2 to 10 atm. 3. The method according to claim 2, wherein said pressure is about 3-5 atm. 4. The method according to claim 1, wherein cumyl hydroperoxide is cracked in a multi-reactor, said multi-reactor is several tubular circulation reactors with shells connected in series, and dicumyl hydroperoxide Alkenyl peroxides are cracked in a unidirectional flow reactor. 5. A process according to claim 4, wherein cumenyl hydroperoxide is cracked in three reactors at temperatures ranging from: First Reactor: From about 67°C to about 82°C. Second Reactor: From about 63°C to about 82°C. Third Reactor: From about 57°C to about 70°C. 6. A method according to claim 5, characterized in that said temperature range is: First Reactor: From about 68°C to about 79°C. Second Reactor: From about 65°C to about 78°C. Third Reactor: From about 60°C to about 69°C. 7. The process according to claim 1, wherein after the second reactor, the pressure is reduced to about atmospheric pressure or lower, and the acetone present in the lysate is partially evaporated, concentrated and returned to the first reactor , and before the acetone portion is evaporated but after the lysate leaves the second reactor, ammonia is added to the lysate, thus turning the remaining acid catalyst into a neutral salt. 8. The method according to claim 1, characterized in that technical grade cumenyl hydroperoxide feedstock contains about 74%wt to about 92%wt cumenyl hydroperoxide, and is recycled to cumenyl hydroperoxide cracking Acetone in the reactor is calculated according to the following formula: C _ac =G _CHP) ×0.125[CHP]+ 35/(G _CHP [CHP]) Wherein: Gac is the amount of acetone added for cracking, t/hr, G _CHP is the amount of technical grade cumyl excess hydrogen oxide added for cracking, t/hr, and [CHP] is the weight concentration of cumenyl hydroperoxide in technical grade cumenyl hydroperoxide. 9, according to the method for claim 1, wherein in the cumenyl hydroperoxide cracking product, the cumenyl hydroperoxide containing 40-98%wt and the acetone that circulates in the cumenyl hydroperoxide cracking reactor, according to Calculate the following formula: G _ac ＝0.58G _CHP × ((1.315[CHP])/(50+0.25[CHP]) - (2[CHP])/152 - ([DMBA])/136 - ([cumene])/120 - ([AP])/120 ) Wherein: Gac is the amount of acetone added for cracking, t/hr, G _CHP is the amount of technical grade cumenyl hydroperoxide added for cracking, t/hr, [CHP] is the weight concentration of cumenyl hydroperoxide in technical grade cumenyl hydroperoxide. [DMBA] is the weight concentration of dimethyl benzyl alcohol in technical grade cumenyl hydroperoxide [Cumene] is the weight concentration of cumene in technical grade cumyl hydroperoxide [AP] is the weight concentration of acetophenone in technical grade cumenyl hydroperoxide. 10, according to the method for claim 1, wherein a part of cumenyl hydroperoxide cracking product is circulated to cumenyl hydroperoxide reactor by circulation pipeline, and the first calorimeter of inlet and outlet is connected in parallel to above-mentioned circulation pipe On the way, a part of the circulating cumenyl hydroperoxide cracking product passes through the above-mentioned calorimeter, and the temperature of this part is measured at the inlet and outlet of the above-mentioned first calorimeter to obtain the first temperature difference, and the second A calorimeter with an inlet and an outlet is connected in parallel to the pipeline connecting the cumenyl hydroperoxide cracking reactor and the dicumenyl peroxide catalytic cracking reactor, and is in the state where water is added to the cumenyl hydroperoxide At the position after the cleavage product, a part of the cumenyl hydroperoxide cleavage product sent to the dicumenyl peroxide cleavage reactor passes through the inlet and outlet of the above-mentioned second calorimeter, and in this second calorimeter Measure the temperature of this part at the inlet and outlet of the device to obtain the second temperature difference, the absolute difference between the first and second temperature difference, that is, δT=δT ₂ -δT ₁ , the range is about 0.2°C- within 3°C. 11. The process according to claim 1, wherein a portion of the cumyl hydroperoxide lysate is recycled through the first reactor in a recycle ratio of about 20-40 parts of the recycled lysate to one part of the added cumenyl hydroperoxide. hydrogen peroxide. 12. A process according to claim 11 wherein said ratio is about 26-35 parts by weight of recycled lysate to one part by weight of added cumenyl hydroperoxide.

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