MXPA00004379A - High selective method of phenol and acetone production - Google Patents

Abstract

Disclosed is a process for the cleavage of technical cumene hydroperoxide (CHP) into phenol, acetone and&agr;-methylstyrene. In a first stage, the CHP cleavage process is conducted in such a way to maintain the heat generation rate and the heat removal rate balanced in each of the CHP cleavage reactors. The cleavage of the CHP is conducted under substantially isothermal conditions at a temperature in the range of 47-50°C. In the second stage of the process dicumylperoxide (DCP) and dimethylbenzene alcohol (DMBA) cleavage is carried out in a multi-section plug-flow reactor under non-isothermal conditions at a controlled temperature increase. The temperature is controlled with the use of thermocouples installed in each section of the reactor. The obtained temperature profile is compared with the temperature profile required by the kinetic model based on&Dgr;T in each section of the reactor. Based on any obtained fluctuations at least one of the amount of water additionally fed to the reactor, the temperature and the degree of sulfuric acid conversion into NH4HSO4 are adjusted.

Description

HIGHLY SELECTIVE METHOD FOR THE PRODUCTION OF PHENOL AND ACETONE BACKGROUND OF THE INVENTION The present invention relates to the field of petrochemical synthesis and in particular, to a method for the production of phenol, acetone and alpha-methylstyrene (AMS) by the cumene method.

There are several known methods for producing phenol and acetone by using the acid cleavage of technical cumene hydroperoxide (CHP). The main difference between the known methods is the use of different reaction media and alternative techniques to remove the heat (380 Kcal / kg) generated during the process of splitting the CHP.

In the processes of these prior techniques the best selectivity is obtained with the use of an equimolar mixture of phenol and acetone as the reaction medium. On a relative basis, 15-30% acetone, based on technical CHP, is added to this mixture. This is illustrated in the Russian Application No. 9400736/04/007229 dated lo. March 1994 and US Patent No. 4,358,618. This allows obtaining a good selectivity of the process that is determined by the REF. 120102 Obtaining the desired by-product, AMS, formed from the alcohol dimethylbenzene (DMBA) present in the technical cumene hydroperoxide. The yield obtained from AMS is 80%.

During the splitting of the CHP, the generated heat is removed. In the process according to US Pat. No. 2,663,735 the heat is removed by the evaporation of acetone and the recirculation of acetone to the reactor. The heat generated can also be removed by the use of a cooling medium such as cooling water.

During the adiabatic cleavage of 100% CHP the temperature increases to about 700 ° C under the influence of an acid catalyst. The heat is generated spontaneously. Due to the rapid release of heat, the process of splitting the CHP is considered very dangerous. Consequently, the combination of heat generation and heat removal is a high priority for the improvement of process safety.

In the process of US Patent No. 2,663,735, the heat of reaction is removed by evaporation of acetone and heat generation and heat removal are completely combined. The heat generated in the process for the splitting of 1 ton of CHP requires the feeding of approximately 2.2-3 ton of acetone to the reactor. The evaporated acetone is extracted from the reactor, condensed and continuously recycled to the reactor. As a result, the reactor is operated in a thermally stable manner as required for the safety of the process.

However, the stable thermal condition is obtained only by the use of a comparatively high sulfuric acid concentration of 1200-1300 ppm. However, the high concentration of H SO., Which is needed given the large amount of acetoda fed into the cleavage products, reduces the activity of sulfuric acid which is the CHP splitting catalyst. Therefore, a high concentration results in a low yield of the desired products and a high content of microimpurities (about 1500 ppm) such as mesityl oxide, hydroxyacetone, and 2-methylbenzofuran which substantially adulterate the quality of the phenol. While the chemistry of the process requires a low concentration of sulfuric acid of about 100-300 ppm, this in practice can not be carried out since the CHP accumulates in the bottoms of the reactor due to the sharp decrease of the CHP rupture that as a result it gives a great heat release, that is, when the concentration of sulfuric acid is reduced the reactor is operated under unstable thermal conditions. Actually the process reaches the thermal stability only at a high concentration of sulfuric acid but this results in a low selectivity of the process. Therefore, in the process that uses acetone evaporation, the objective of thermal stability and obtaining a high selectivity are an irreconcilable conflict.

In the process of the above reference to the Russian application, US Patent No. 4,358,618 and US Patent No. 5,254,751, the heat of reaction is removed with the reaction products or the mass, from the splitting reaction (RCM) by means of multiple circulations through the water-cooled heat exchangers. The heat exchangers, which can be named from 2 to 6, are in fact the reactors in which the splitting of the CHP occurs. The thermal stability of the process (ie the safety of the process) depends on the composition of the reaction products, the range of the acid concentration, the temperature profile and, therefore, the distribution of the CHP conversion in the reactors The stability of the process deteriorates to a greater conversion of CHP in the first reactor and when the temperature difference between the first and the subsequent reactors increases. In practice, the more non-isothermal the conditions of the process, the more precarious is the state of the process.

In the process in accordance with the Russian application, the splitting of the CHP and the DCP are carried out in two stages. The splitting reactors of the CHP (mixing reactors) and the conversion of the DCP (piston flow reactor) are operated at the same pressure.

The splitting of the CHP and the DCP is carried out in an equimolar mixture of phenol and acetone containing up to 12% weight of cumene. To reduce the acidic properties of the sulfuric acid and, therefore, to increase the yield of such desired products such as phenol, acetone and AMS, additional acetone is added in the reaction products according to the following algorithm: Gac _ GCHP x 0. 125 [CHP] + 35 / (Gr CHHPP x [CHP]) where: Gac, GCHp represent the flow velocity of acetone additional technical CHP, respectively, in kg / hr and [CHP] is the technical grade CHP concentration (% by weight) which is equal to 12-14% relative to acetone based on the technical, CHP feed rate.

The conversion of CHP, depending on the feed rate, is maintained in the first reactor of 62-75%, in the second reactor of 87-94% and in the third reactor of 94-98%. The corresponding temperatures in these reactors are 67-79 ° C, 78-67 ° C and 69-60 ° C, respectively. The above algorithm for the additional acetone feed, the temperature, and the CHP conversion distribution in the reactors allow the process to operate over a wide range of reaction rates.

The concentration of CHP in the output of the reactors of the first stage is 0.14-0.43% by weight, which corresponds to a? T of 1-3 ° C in the calorimeter that controls the first stage of the process.

The water is fed to the splitting reactor of the DCP in an amount such that it provides a water concentration in the reaction products of 1.3-2.0% by weight. The operation of the reactor of the second stage is controlled by a? T equal to 1-3 ° C of the calorimeter installed in the line before the splitter reactor of the DCP. In the splitter reactor of the DCP the process conditions are isothermal. In the DCP splitting reactor, different temperatures of 94 ° C are maintained at low feed rates at 99 ° C at high feed rates. The whole process (the. And 2nd stages) is controlled by the temperature differential between the two calorimeters. This temperature differential of the calorimeter? is 0.2-0.3 ° C.

In order to reduce the non-selective losses in the acetone flash stage, it is added in the line before the evaporator to convert the sulfuric acid into the neutral salt (NH4) 2S04. As a result, the theoretical yield of 78.8-79.6% of AMS is obtained in the process.

BRIEF DESCRIPTION OF THE INVENTION An object of the present invention is to provide a process for obtaining a higher yield of desired products by increasing the yield of AMS to 85-87% and reducing the chemical losses in the rectification columns of the rupture product.

Another objective is to increase the safety of the process by splitting the CHP under conditions that are substantially isothermal.

Further objectives of the invention are the reduction of the energy consumption in the process by the reduction of the amount of recirculating acetone and the recovery of the heat with the splitting reactor of the DCP and DMBA and to obtain a stable conversion of DCP in the second stage of the process at variable feed rates and fluctuating operating conditions.

It is a further objective of the invention to reduce non-selective losses in the rectification stage of the splitting product. These objectives and others are obtained by the process of the invention.

In the process of the invention, the technical grade CHP containing DMBA is split into phenol, acetone and α-methylstyrene.

The CHP technical grade is introduced into at least the first of a series of at least three sequential reactors in which the CHP is unfolded under the influence of an acid catalyst. The reactors are maintained under substantially isothermal conditions in a temperature range of about 47 to 50 ° C to produce a product stream containing DCP and DMBA. The product stream is introduced into a splitting reactor in which the DCP is decomposed in a non-isothermal operation in a mixture containing at least phenol, acetone and α-methylstyrene.

The advantages of the invention are obtained by the selection and control of the temperature conditions in the first and second splittings, by the conversion of the CHP in the reactors of the first stage, by the composition of the reaction products and by the change of the algorithm of the control of the reactor in the second stage of the process.

The process of the invention, similarly, to known processes of phenol, comprises several main steps that determine the selectivity of the process as a whole: 1. Oxidation of cumene (isopropylbenzene) with air and / or oxygen to cumene hydroperoxide (CHP); 2. Acid cleavage (H2S04) of the CHP produced; and 3. Rectification of the splitting products of the CHP by the multistage rectification method.

The process of the invention shows an improvement of the consumption parameters of the process such as the value of the power consumption. More specifically, it has an improvement in the safety of the splitting of the CHP by the balance in heat generation and heat removal rates and a reduction in steam consumption. The process involves a new principle of control over the second stage - conversion of dicumylperoxide (DCP) and alcohol dimethylbenzene (DMBA). A decrease in the chemical losses of the desired products is shown in the rectification stage obtained by the change in the composition of the products at the outlet of the DCP reactor.

The various novel characteristics characterizing the invention are pointed out with particularity in the appended claims and which form part of this specification. For a better understanding of the invention, its operational advantages and specific objectives obtained by its use, reference should be made to the accompanying drawings and descriptive material in which a preferred embodiment of the invention is illustrated and described.

BRIEF DESCRIPTION OF THE DRAWINGS In the drawings, in which similar reference characters denote corresponding or similar elements through the various figures: Figure 1 shows schematically one embodiment of the process of the invention; Figure 2 shows the profile change in the DCP reactor depending on the amount of heat generated by the splitting products; and Figures 3A and 3B respectively show the dependence of the temperature profile on the conditions of the DCP reactor and the dependence of the concentration of DCP on the conditions of the DCP reactor.

DETAILED DESCRIPTION OF THE INVENTION The process of unfolding of CHP and DCP of the invention can be seen having a first and second stage for the purposes of the description. In the first stage, the CHP is split and the DCP is synthesized in mixing reactors. This splitting is conducted under the influence of an acid catalyst which is preferably sulfuric acid.

Referring to Figure 1, a technical CHP feed stream 10 or containing cumene oxidized to CHP in accordance with processes of the prior art and containing DMBA is introduced into a first of a cascade of reactors 12. In a preferred embodiment, the cascade 12 includes three reactors 14, 16 and 18 arranged in series. The reactors 14, 16 and 18 are mixing reactors with respect to the reactions of by-products, and piston-flow reactors with respect to the decomposition reactions of the CHP. For this purpose, a series of deflectors (not shown) are installed in a part of the body of each of the reactors 14, 16 and 18 to allow the conversion of the reactors 14, 16 and 18 of the mixing regime to the flow rate in piston in each section of each reactor. Preferably, six to six deflectors are installed in each reactor.

In reactors 14, 16 and 18 the CHP is unfolded to form a first product stream 20 containing 1% CHP, phenol and acetone, 4-5% DCP at 2-2.5% DMBA, approximately 1-1.5% AMS, and minimum amounts of byproducts - dimers of AMS and complex phenols. The splitting is carried out by the sulfuric acid which is at a concentration in the reaction products of not less than 180 ppm and not higher than 200 ppm. The first product stream 20 leaving reactor 18 is divided and a portion of that stream is recycled through line 22 to a pump 24 from where the material is sent to reactor 14 after being combined with the CHP grade feed stream 10. The relative amount of the fraction recycled from stream 20 to stream 10 is about (8-40): 1. The sulfuric acid 26 can be introduced into the recycle line 22.

Additional acetone is fed in the first stage of the process. The amount of additional acetone is based on the flow rate of the technical CHP and is maintained in the range of 5-8% relative to the flow rate of CHP to achieve the required CHP conversion value at variable feed rates and fluctuating operating conditions. The amount of additional acetone fed should not exceed 8%.

The conversion of CHP in reactors 14, 16 and 18 in series is maintained at 42-50%, 67-73%, and 78-82%, correspondingly. The temperature in each of the reactors 14, 16 and 18 is maintained between 47 and 50 °. The cooling water removes the heat generated in the process. Preferably the temperature in reactor 14 is 47-50 ° C, in reactor 16 it is 50-48 ° C, and in the reactor it is 48-50 ° C. That is, unlike the prior art as illustrated in the Russian application cited above and in U.S. Patent No. 5, 254, 751, the process conditions in the process of the invention are isothermal or at least substantially isothermal in the reactors 14, 16 and 18. The previous distribution of the CHP conversion and the temperature in the reactors allow to balance the heat generation speed and the heat removal speed by controlling the CHP splitting speed. This balance results in a system in which heat stabilizes at all points of the reactors, thus promoting the safety of the process.

A substantially isothermal operation in reactors 14, 16 and 18 is obtained by operating with certain amounts of additional acetone fed, certain concentrations of water in the reaction products and obtaining a lower concentration of acid in the reaction products. The combination of the above aspects results in a certain rate of splitting of CHP and as a result a certain amount of heat in each of the reactors 14, 16 and 18 in the first stage. Due to different flow rates of cooling water to the reactors of the first stage, the isothermal or closely isothermal conditions are maintained. When? Tl - the difference between the outlet and inlet temperatures of a flow calorimeter 28 - deviates from the required temperature by 8-9 ° C, the temperature is corrected in the first reactor which maintains the required CHP conversion value, and, as a result, the temperature after the first reactor. The temperature after the last reactor of the first stage is also maintained by controlling the flow rate of the cooling water. Also the cooling water is fed into the tube space of the second reactor but that flow rate is preferably kept stable at constant CHP feed rates. Such an operation method provides conditions that are isothermal or substantially isothermal in the CHP splitting reactors in the first stage.

An advantageous aspect of the process of the invention is that it eliminates the areas of temperature increase in the reactor that occur in the splitting methods of the conventional prior art. The rate of formation of undesirable byproducts, such as AMS dimers and complex phenols, is reduced resulting in an increase in the selectivity of the CHP cleavage step, and, as a result, the overall selectivity of the process.

The system to carry out the process includes a temperature measurement arrangement that in Figure 1 is illustrated as the calorimeter 28.

The remaining non-recycled portion 30 of the product stream 22 is introduced into an intermediate container 32. The water 34 and a base 36 which is preferably NH 4 OH are mixed with the product stream 30 in the container 32. As indicated, a Temperature measurement is carried out by means of a temperature measuring device 40 shown as a calorimeter in the discharge line of tank 38. The mixed stream in line 38 is preferably heated in two stages by heat exchangers 42 (from 80- 90 ° C) and 44 (90-100 ° C) in such a way that the temperature of the current increases by around 50-55 ° C.

The heated stream 46 is introduced into a reactor 48 for the splitting of the DCP and the dehydration of the DMBA. The reactor 48 is preferably a multistage piston-flow reactor with an internal arrangement of deflectors that form a plurality of sections or zones within the reactor 48.

In the piston flow reactor 48, the main reaction of the conversion of DCP to phenol, acetone and AMS and the reaction of lateral conversion of DMBA to the desired by-product AMS takes place. AMS is a desired product since it can be converted into cumene and then returned to the cumene oxidation stage.

In reactor 48, the feed temperature is raised in a controlled manner to a temperature in the range of about 120 to 150 ° C and preferably 140-146 ° C. The change in reactor 48 is a self-maintenance reaction. Preferably, each section or zone of the reactor 48 has an independent temperature control by means of, for example, a thermocouple and feedback temperature control and subsequent systems.

A product stream 50 leaves the reactor 48, passes through the heat exchanger 42, where it transfers heat to the stream 38 and enters the evaporator 56 where the evaporation of the additional acetone fed takes place. The stream 52 exits through the heat exchanger 42 and mixes with a base, such as NHOH (54), and then passes to an evaporator 56 where part of the acetone is evaporated together with parts of water, cumene and phenol. The evaporated phase 58 is condensed in the condenser 60, separated and the condensed acetone 66 is recycled. A portion 68 of the recycled acetone 66 is introduced into the intermediate vessel 32 while the portion 70 is mixed with the stream 22. The non-evaporated splitting products 72 are removed from the evaporator 56 and cooled in the heat exchanger 74 and exit as stream 76. For the acetone added in the first and second steps, crude acetone can be used for the final products of the distillation step of the acetone columns (not shown).

In the first and second stages of the process along with the desired process products such as phenol, acetone and AMS, undesirable byproducts such as dimers of AMS and complex phenols are formed in the reactors.

The formation of by-products is considered to take place by means of the mechanism of the carbon and the conventional ion of the acidic and catalytic reactions, that is, the products are protonated by means of the double ligation of the AMS to form carbocationite "A" H + + C6H5- -C = CH0 = C6H5"-C-CH. (A) CH3 CH and the additional conversion of "A" into complex phenols and AMS dimers.

C6H5OH * - ortho and para-cumyl phenol dimers C6H5 C CH2 + CH "+ C6Hs C - CH2 •" AMS CH3 However, it has now been discovered that the reactive particle is not a carbon ion but an oxonium ion (B) is formed: CH, CH, H + C'fif6H '"5? -C -OH C6H-C -OH2 CH, CH, When phenol and DMBA react with this oxidation, AMS dimers and complex phenols are formed. Therefore the reactive particle is not AMS but a molecule of DMBA.

The determined reaction mechanism required further investigations of the conversion reaction conditions of DMBA to AMS and DCP.

In fact, the balance of the process recovers: CH3 C6H5- C = CH2 + H30 + Two important factors, such as solvent composition (ie, product composition with respect to the process) and temperature, simultaneously influence the balance between the first and second reactive particles. Having determined its reaction mechanism, the reaction conditions of the DMBA conversion in the DCP reactor were reexamined.

The displacement of the previous equilibrium results in three to four times the decrease in the unreacted amount of DMBA and the formation of undesirable products. It also results in an AMS yield of theoretical 85-89.7% under the selected DMBA and DCP conversion conditions in the second stage of the process. In addition, the decrease in the DMBA content at the outlet of the DCP reactor results in a decrease in the amount of undesirable products formed in the distillation columns from about 15-17 kg / t of phenol to about 8-10 kg / t of phenol which is equal to the decrease in the cumene consumption coefficient of about 7-8 kg / t which is equivalent to the economization of the initial cumene product of about 80,000 kg per year for every 100,000 tons of the final phenol product. The approach described above of the displacement of the equilibrium in the direction of the AMS allows the reduction of chemical losses during the distillation stage and also allows the increase in the total selectivity of the process, in particular due to the reduction of non-selective chemical losses during the stage of distillation.

Our studies show that the DMBA balance? AMS is established very quickly. Simultaneously, the DMBA reacts to form AMS dimers and complex phenols. The rate of formation of the complex dimers and phenols is slower than the first reaction but increases substantially when the temperature is increased. Therefore, there is an adverse competition between these reactions because when the temperature is increased the equilibrium is shifted to the desired AMS product but also the amount of undesirable products, such as dimers of AMS and complex phenols, is increased. In order to minimize the formation of complex dimers and phenols while improving the performance of AMS, the splitting process in the DCP reactor is conducted in such a way that the average reaction temperature of the DCP reactor is not kept close to the temperature but is preferably maintained in such a way that the average temperature in the reactor or in several zones is lower, for example, preferably about 15 ° C lower, than the maximum temperature reached in the reactor. On the other hand, the temperature should not rise in the respective zones at a very low speed because this also interferes with the production of the desired final products. The real temperature curve depends on the amount of DCP and DMBA that is a function of the selectivity of the first stage, high concentrations of DCP displace the curve.

It has also been found that the effect of the splitting heat of the DCP is 214 kcal / kg. Using the determined heat release of the reaction, the splitting process of DCP is conducted under non-isothermal conditions as shown in Figure 2.

Depending on the amount of heat generated by the splitting products in the heat exchanger 44 (see Figure 1), the temperature profile in the DCP reactor can be different, ie essentially isothermal (curve T-l) is non-isothermal (curve T-3) or an intermediate profile (curve T-2) as shown in Figure 2.

Despite equal temperatures at the entrance and exit of the reactor in the case of Tl and T-2 (Figure 2) and in the case of T-2 and T-3 when the average temperature in the reactor is the same, the results The final yields of AMS are significantly different. The results are obtained for the case T-1 when the temperature in the reactor 48 is almost constant (that is, the conditions are isothermal). Under these conditions the AMS yield is 70% theoretical.

The best results are obtained when the process operates non-isothermally (see curve T-2) in a plug flow reactor 48. The AMS yield is 89.7% theoretical. For the case of T-3 when the average temperature is equal to the average temperature T-2, intermediate results are obtained between the isothermal and non-isothermal processes: the AMS yield is 78-80% theoretical.

In a preferred embodiment, a thermocouple is installed in each section of the DCP reactor to maintain the maximum AMS yield in the DCP reactor. The temperature profile obtained is compared with the optimum temperature profile, the latter based on the developed kinetic model.

In the case of fluctuations in the temperature profile when the DCP conversion is incomplete or the DCP conversion exceeds the allowed value, the water concentration of the reactor is adjusted to return the temperature profile to the initial values, as shown in Figure 3A. Figure 3A describes the dependence of the temperature profile on the conditions of the DCP reactor. In Figures 3A and 3B, curve 1 indicates the optimum temperature profile, curve 2 indicates the profile under severe conditions, and curve 3 indicates in Figure 3A the profile during mild conditions and indicates in Figure 3B the profile when DCO becomes incomplete. Zone "I" indicates the heating zone of the splitting product.

Under the severe conditions of curve 2, Figure 3B, additional water is fed to the reactor. This decreases the acid properties of the catalyst and optimizes the temperature profile.

In the case of incomplete DCP conversion (curve 3, Figure 3B) in the reactor, the amount of water fed to the reactor decreases and the temperature in the heater installed before the DCP reactor increases. This results in an increase in the speed of DCP splitting and allows to obtain the required DCP conversion value.

The process of the invention shows numerous advantages over the process of the prior art. In particular, the inventive process differs from the process described in U.S. Patent No. 5-254-751 as follows: 1. The CHP splitting process in the mixing reactors is conducted, due to the balanced generation of heat and the heat removal rates, that is to say under conditions that are very close to, or are substantially isothermal. Estro improves process safety and selectivity.

The splitting process of DCP in the piston flow reactor is non-isothermal at a controlled temperature increase of 120 to 146 ° C and the depth of conversion of DCP and DMBA controlled by the change of water concentration at the same time in the cleavage products and to the degree of the conversion of sulfuric acid to NH4HS04 at variable flow rates. The temperature is controlled by the installation of a thermocouple in each section of the piston flow. The temperature profile obtained is compared with the temperature profile required by the kinetic model based on the deviation of the temperature or the value? in each section of the reactor and in the fluctuations the amount of water fed additionally to the reactor, the temperature and the degree of conversion of sulfuric acid to NHHS04 are corrected.

The composition of the reaction environment in the decomposition step of the CHP and in the decomposition step of the DCP is materially different due to the addition of varying amounts of acetone in each of the mentioned steps. 4. Due to the change in the composition of the reaction medium and the change of the control algorithm of the reactor in the second stage of the process, the yield of the desired AMS product is increased to theoretical 85-89.7%.

The above advantages and characteristics of the process of the invention are demonstrated by means of the following examples and are tabulated in tables 1 and 2 shown below after the descriptions of the examples. Example 1 is a comparative example while examples 2 to 11 are of the process of the invention.

EXAMPLE 1 (comparative) 72 t / hr of CHO of technical grade are fed to the reactor block comprising three tube-type reactors installed in series. The reactors are operated at pressures of 2-10 atm. The composition of the CHP of technical grade introduced in the series of reactors is the following: Component% by weight Cumene hydroperoxide 82.9 Cumene 12.0 DMBA 4.2 Acetophenone 0.6 DicumiIperoxide 0.3 9976 kg of acetone are added continuously to the cleavage products that circulate according to the algorithm (application 12, 16% of the amount of CHP added).

As a result of the addition of acetone, the molar ratio of the phenol: acetone: cumene reaction products is 1: 1.42: 0.22.

Sulfuric acid is added continuously to the cleavage products that circulate. The flow velocity of sulfuric acid is 21 kg / hr, the content of sulfuric acid in the reaction products is 250 ppm, and the water flow rate is 2 kg / hr.

In the first stage, CHP conversion is maintained at 65% in the first reactor, 89.6% in the second reactor and 94.5% in the third reactor. The temperatures in the respective reactors are maintained at 75.8 ° C, 72.4 ° C and 63.1 ° C.

The concentration of CHP at the outlet of the first stage of reactors (14, 16, 18) is 0.21 by weight - percentage that corresponds to a value of? Tl of 1.59 ° C in the calorimeter through which the first stage of the process It is controlled.

The splitting of the DCP formed in the circulation link is conducted in a two-section adiabatic piston flow reactor. Water at a flow rate of 716.6 kg / hr is continuously added to the feed line to the splitting reactor of the DCP to maintain the water concentration at 1.91% by weight at the outlet of the reactor.

In the DCP splitting reactor the same composition of the reaction products is maintained, ie the phenol: acetone: cumene ratio as found in the CHP cleavage products.

The reactor of the second stage is controlled by a temperature differential? T2 which is equal to 1.34 ° C by the calorimeter installed in the line before the splitting reactor of DCP. The process in the splitting reactor of DCP is isothermal at a temperature of 99 ° C. The overall process (first and second stages) is controlled by the temperature difference between the two calorimeters. That temperature difference based on the readings of the calorimeters at 0.25 ° C.

The acetone added to the reaction products in the unfolding stage of CHP is removed in the evaporator installed after the DCP splitting reactor. After being distilled in the evaporator and condensed in the condenser, the acetone is recycled to the splitting stage of CHP: To decrease the non-selective loss of desired products such as phenol and AMS, aqueous ammonium solution is added to the evaporator to convert the sulfuric acid to the neutral salt (NH4) 2S0.

The conversion of AMS after the splitting stage is 78.6% theoretical.

EXAMPLE 2 72 t / hr of technical grade CHP having the composition of Example 1 are introduced into the mixing reactors in the CHP splitting step. The process is as shown in Figure 1.

The splitting of CHP is indicated in the reaction products where the molar ratio of phenol: acetone: cumene is maintained at 1: 1.28: 0.22. This corresponds to an 8% acetone additionally fed based on technical CHP.

The flow rate of sulfuric acid is 16.6 kg / hr. The concentration of sulfuric acid in the reaction products is 200 ppm.

In the first reactor, a CHP conversion of 50% is maintained. In the second reactor, the conversion is 69.0% and 81.6% in the third reactor. The temperatures are 48.2 °, 48.3 ° and 49 ° C, respectively. The temperature profile in the CHO splitting reactors is substantially isothermal.

The splitting of DCP is conducted in a multisection piston flow reactor that operates non-isothermally at a controlled temperature increase of 120 to 137 ° C.

Each section of the reactor is equipped with a system to maintain a set temperature there.

Water at a flow rate of 418.9 kg / hr is continuously added to the feed line in the DCP splitting reactor to maintain the water concentration at 1.4% by weight at the outlet of the reactor. 57.5 kg / hr of a 5% aqueous solution of ammonium is added to achieve a conversion of NH4HSO4 of 50%.

The acetone, added to the reaction products of the CHP cleavage stage, is removed in the evaporator which follows the DCP splitting reactor. The acetone, distilled in the evaporator and condensed in the condenser, is recycled to the CHP splitting stage. To reduce the non-selective losses of the desired products such as phenol and AMS, an ammonia solution is added to the evaporator to convert sulfuric acid to the neutral salt (NH) 2SO, |. The performance of AMS after the splitting stage is 85.6% theoretical.

EXAMPLE 3 A CHP split process is conducted in the same manner as in Example 2 with CHP of technical grade of the following composition: Component% by weight Cumene hydroperoxide 90. . 3 Cumeno 2. . 0 DMBA 6. . 2 Acetophenone 1,. 0 Dicumiperoperoxide 0. . 5 The conversion of CHP in the first reactor is' 49.6%, in the second reactor it is 67.0% and in the third reactor it is 78.9%. The temperatures in the reactors are 48.5 ° C, 49.5 ° C and 50.0 ° C, respectively.

The splitting of DCP is carried out in a multisection piston flow reactor that operates non-isothermally with a controlled temperature increase of from 120 to 143 ° C equipped with a forced temperature independent maintenance system fixed in each section.

A continuous water flow of 198.7 kg / hr is added to the feed line in the DCP splitting reactor to maintain the water concentration at 1.4% at the outlet. An aqueous solution of 5% ammonium is added at a flow rate of 57.5 kg / hr to achieve a conversion of sulfuric acid to NH4HS04 of 50%.

The AMS yield reached after the splitting stage is 85.1% theoretical.

EXAMPLE 4 The process is conducted in the same way as in the Example 2 except that 15.1 kg / hr of sulfuric acid are added to the cleavage products that circulate which produces a decrease in the concentration of H2SO4 in the CHP splitting reactors at 180 ppm.

The conversion of CHP in the first reactor is 48.8%, in the second reactor it is 67.0% and in the third reactor it is 79.6%. The temperatures are 48.4 ° C, 49.1 ° C and 49.9 ° C, respectively.

The splitting of DCP is carried out in a multisection piston flow reactor operating non-isothermally at a controlled temperature increase of from 120 to 139 ° C. The reactor is equipped with a system to independently maintain the fixed temperature in each section.

The conversion of AMS after the splitting stage is 85.8% theoretical.

EXAMPLE 5 A CHP cleavage process is conducted in the same manner as in Example 2 except that the cleavage is conducted in the reaction products maintaining the molar ratio of phenol: acetone: cumene in 1: 1.19: 0.22 corresponding to a 5% relative acetone fed additionally based on CHP of technical grade.

The concentration of H2SO4 in the reaction products is 180 ppm.

The conversion of CHP in the first reactor is 50.0%, in the second reactor it is 68.8% and in the third reactor it is 81.7%. The temperatures are 47.0 ° C, 48.3 ° C and 48.9 ° C, respectively.

The splitting of DCP is carried out in a multisection piston flow reactor that operates non-isothermally at a controlled temperature increase from 120 to 135 ° C. The reactor is equipped with a system to independently maintain the fixed temperature in each section.

The conversion of AMS after the splitting stage is 85.7% theoretical.

EXAMPLE 6 A CHP split process is conducted in the same manner as in Example 4 except that the feed rate is 90 t / hr, that is, 25% higher than in Comparative Example 1.

The CHP conversion in the first reactor is 44.0%, in the second reactor it is 67.0% and in the third reactor it is 77.1%. The temperatures are 50.0 ° C, 50.0 ° C and 48.6 ° C, respectively.

The splitting of DCP is carried out in a multisection piston flow reactor operating non-isothermally at a controlled temperature increase of from 120 to 137 ° C. The reactor is equipped with a system to independently maintain the fixed temperature in each section.

The conversion of AMS after the splitting stage is 85.6% theoretical.

EXAMPLE 7 A CHP splitting process is conducted in the same manner as in Example 4 except that the feed rate is 54 t / hr, that is, 25% lower than in Comparative Example 1.

The conversion of CHP in the first reactor is 50.0%, in the second reactor it is 72.09% and in the third reactor it is 81.9%. The temperatures are 50.0 ° C, 49.2 ° C and 49.0 ° C, respectively.

The splitting of DCP is carried out in a multisection piston flow reactor operating non-isothermally at a controlled temperature increase of from 120 to 137 ° C. The reactor is equipped with a system to independently maintain the fixed temperature in each section.

The conversion of AMS after the splitting stage is 85.5% theoretical.

EXAMPLE 8 A CHP splitting process is conducted in the same manner as in Example 4 except that water at a flow rate of 886.0 kg / hr is added to the CHP cleavage products before they are fed to the splitting reactor. of DCP to maintain the water concentration equal to 2.0% by weight in the DCP splitting reactor.

The splitting of DCP is performed non-isothermally at a controlled temperature increase of from 129 to 146 ° C.

The conversion of AMS after the splitting stage is 87.0% theoretical.

EXAMPLE 9 The CHP splitting process is conducted in the same manner as in Example 7 except that water at a flow rate of 629.0 kg / hr is added to the CHP cleavage products before they are fed to the splitting reactor. of DCP to maintain the water concentration equal to 1.7% by weight in the DCP splitting reactor.

The splitting of DCP is performed non-isothermally at a controlled temperature increase from 125 to 142 ° C.

The conversion of AMS after the splitting stage is 86.4% theoretical.

EXAMPLE 10 A CHP splitting process is conducted in the same manner as in Example 2 except that water at a flow rate of 886.0 kg / hr is added to the CHP cleavage products before they are fed to the splitting reactor. of DCP to maintain the water concentration equal to 2.0% by weight in the DCP splitting reactor. The splitting stage of DCP is performed non-isothermally at a controlled temperature increase of from 129 to 146 ° C.

The conversion of AMS after the splitting stage is 86.8% theoretical.

EXAMPLE 11 A process of unfolding of CHP is conducted in the same manner as in Example 9 except that water at a flow rate of 629.0 kg / hr and 17280 kg / hr of acetone is added to the CHP cleavage products before the feed is introduced into the feed. DCP splitting reactor to maintain the water concentration equal to 1.7% by weight in the splitting reactor of DCP and the additional concentration of acetone in the splitting reactor of DCP in 24% relative to the CHP introduced in the first stage.

The splitting of DCP is performed non-isothermally at a controlled temperature increase from 125 to 142 ° C.

The conversion of AMS after the splitting stage is 89.7% theoretical.

The results of the previous examples are tabulated in the following Tables 1 and 2.

Summary of Examples Table 1. la. Stage - unfolding of CHP Summary of Examples Table 1. la. Stage - unfolding of CHP (continued) Summary of Examples Table 2. 2a. Stage - Split of DCP The terms and expressions that have been used and used as terms of description and not limitation, and there is no intention for the use of such terms to exclude any equivalence of the characteristics shown and described or their portions, it is recognized that several modifications are possible in the scope of the invention.

It is noted that in relation to this date, the best known method for carrying out the aforementioned invention is that which is clear from the present description of the invention.

Having described the invention as above, the content of the following is claimed as property.

Claims (11)

1. A process for the splitting of technical CHP containing DMBA to phenol, acetone and α-methylstyrene, characterized in that it comprises: the introduction of technical grade CHP in at least the first of a series of at least three sequential reactors; the splitting of CHP in said three sequential reactors under substantially isothermal conditions in a temperature range of about 47-50 ° C to produce a first product stream at a first temperature, said product stream containing DCP; heating said first product stream to a second temperature of 50-55 ° C higher than said first temperature, thereby producing a first stream of heated product; and the introduction of the first product stream in a plug flow reactor where the DCP is decomposed into a mixture containing phenol, acetone and α-methylstyrene.
2. 2. The process of claim 1, characterized in that at least three reactors are operated at a pressure of 1 to 10 atmospheres.
3. 3. The process of claim 1, characterized in that the temperature in the first reactor is in the range of 47-50 ° C, the temperature in the second reactor is 48-50 ° C and the temperature in the third reactor is 48-50 ° C. 50 ° C.
4. 4. The process of claim 1, characterized in that the conversion of CHP in the first reactor is 43-50%, in the second reactor it is 67-73% and in the third reactor it is 78-82%.
5. 5. The process of claim 1, characterized in that the DCP is decomposed in a multisection piston flow reactor.
6. 6. The process of claim 1, characterized in that the cleavage of the CHP is carried out under the influence of an acid catalyst, preferably sulfuric acid, present in an amount of 0.018 to 0.020% by weight.
7. 7. The process of claim 1, characterized in that the splitting of the DCP is conducted under non-isothermal conditions.
8. 8. The process of claim 7, characterized in that the splitting of the DCP is conducted at a temperature of about 120 to about 146 ° C.
9. 9. The process of claim 5, characterized in that the temperature is controlled in each section of the multisection piston flow reactor.
10. 10. The process of claim 8, characterized in that the temperature is controlled by obtaining a temperature profile in each section of the piston flow reactor and comparing the profile obtained with a pre-established temperature profile for the respective section of the reactor.
11. 11. The process of claim 10, characterized in that in response to detected deviations, at least one of an amount of water fed to the piston flow reactor, the temperature, or the degree of conversion of sulfuric acid to NH 4 SO 4 is adjusted. The process of claim 1, characterized in that additional acetone is introduced in an amount of 5 to 8% by weight in relation to the flow rate of the CHO in said three sequential reactors. The process of claim 1, characterized in that additional acetone is introduced in an amount of 8 to 16% by weight, based on the flow rate of the CHO inside the plug flow reactor where the DCP decomposes. A system for the splitting of technical CHP to form phenol, acetone and α-methylstyrene, characterized in that it comprises: a supply line for supplying a feed stream containing CHP; a first stage in which said feed stream is introduced, said first step comprising a series of mixing reactors wherein the CHP is unfolded to produce a product stream; a recycling system by means of which at least a portion of product stream is divided from the remainder of said product stream; a mixing tank for receiving the remainder of said product stream with means for optionally introducing other materials into said tank; at least one heat exchanger for increasing the temperature of the remaining product stream by at least 50-55 ° C; a multistage piston flow reactor within which the stream of heated product is introduced and where the DCP is unfolded; and a temperature measurement and control system adapted to control the temperature increase in each stage of said multistage reactor. A two-stage process for the splitting of technical CHP containing DMBA, phenol, acetone and α-methylstyrene, characterized in that it comprises: the introduction of the technical grade CHP within a first stage, said first stage comprising at least one primer, second and third reactor, said first, second and third reactor being in series; the splitting of CHP in said reactors under substantially isothermal conditions wherein the first reactor operates at a temperature of 47-50 ° C and provides a CHP conversion of 43-50%, the second reactor operates at a temperature of 48-50 ° C C and provides a CHP conversion of 67-73%, and the third reactor operates at a temperature of 48-50 ° C and provides a first product stream at a first temperature, said product stream containing DCP; and introducing a first product stream into a second step, said second step comprising a reactor wherein the DCO is decomposed into a mixture containing phenol, acetone and α-methylstyrene under non-isothermal conditions. The process of claim 15 characterized by the first product stream being heated to a second temperature of about 50-55 ° C higher than the first temperature before being introduced into said second stage. The process of claim 15 characterized in that from 5-8% relative of acetone based on 1 ton of CHP technique is added to the first stage of reactors. The process of claim 17 characterized in that the splitting of DCP is conducted in the second stage in the piston flow reactor and where 8-16% relative of acetone on the basis of 1 ton of technical CHP is added to said reactor . The process of claim 18, characterized in that the weight ratio of additional acetone with r "is similar to the first and second steps is from 1: 1 to 1: 3. The process of claim 20 characterized by the acetone which is added to the reactors of the first and second stages is removed in an evaporator in a vacuum of 200-600 mm Hg, is condensed in a cooler, and is recycled to the regulators. the first and second stages. The process of claim 18 characterized by the crude acetone of the acetone distillation columns is used as additional acetone fed to the first and second stages. The process of claim 15 characterized by the splitting of DCP is conducted non-isothermally under a controlled temperature rise of about 120 ° C to about 146 ° C. The process of claim 15 characterized by the concentration of H2SO4 as a catalyst is 0.018-0.020% by mass in the first stage and 0.005-0.008% by mass maintained in the second stage. The process of claim 15 characterized by the conversion of DCP to the second stage is controlled by a change, optionally simultaneously, of the concentration of water in the reaction medium, of the degree of transfer of sulfuric acid to NHHS04 and of the temperature due to the installation of a thermocouple in each section of the reactor and by comparison of the temperature profile obtained with aguel reguerido by the kinetic model. The process of claim 24 characterized by the profile being controlled in each section of the DCP reactor based on the temperature measurement. The process of claim 25 characterized by the concentration of unreacted DCP after the reactor of the second stage is 0.05 wt% to 0.10 wt%. The process of claim 20 characterized by an aqueous solution of ammonia is added to the evaporator to convert the H2SO to a neutral salt (NH) 2S04 in order to reduce non-selective losses of desired products while evaporating the acetone. The process of claim 1 characterized in that the CHP splitting regime is conducted in accordance with a piston flow reactor regime. The process of claim 28 characterized by facilitating the CHP splitting regime, each of the three sequential reactors comprises a plurality of deflectors. The process of claim 29, characterized in that the plurality of baffles comprises 6-16 deflectors. The process of claim 29 characterized by a magnitude of recirculation ratio of split products with respect to technical CHP is maintained at about 8-40: 1.