WO2019168437A1

WO2019168437A1 - Method for producing olefinic hydrocarbons

Info

Publication number: WO2019168437A1
Application number: PCT/RU2019/000066
Authority: WO
Inventors: Станислав Михайлович КОМАРОВ; Александра Станиславовна ХАРЧЕНКО; Алексей Александрович КРЕЙКЕР
Original assignee: Акционерное общество "Специальное конструкторско-технологическое бюро "Катализатор"
Priority date: 2018-02-27
Filing date: 2019-02-05
Publication date: 2019-09-06
Also published as: RU2655924C1

Abstract

The present method is characterized in that feedstock vapors are heated to a temperature of 150-180°С in a heat exchanger mounted downstream of a recovery boiler (15), wherein dehydrogenation is carried out at a volumetric flow rate of the feedstock of 120-200 h^-1, a temperature of 530-600°С and a linear velocity of contact gas at the outlet from a fluidized bed of 0.5 m/s and higher. The technical result of the proposed invention is to raise the productivity of units for the dehydrogenation of paraffinic hydrocarbons С₃-С₅, to increase the yield of olefinic hydrocarbons and to decrease catalyst consumption.

Description

METHOD FOR PRODUCING OLEFIN HYDROCARBONS

Technical field

The invention relates to the field of petrochemistry, in particular to processes for the production of olefin hydrocarbons used in the production of synthetic rubbers, plastics, high-octane gasoline components and other important organic products.

State of the art

A known method of producing olefin hydrocarbons by dehydrogenation of the corresponding paraffin hydrocarbons in a reactor-regenerator system with a moving coarse-grained catalyst (Y. Ya. Kirnos, O. B. Litvin, “Modern industrial methods for the synthesis of butadiene.” Analytical comparative reviews CNIITENeftekhim, series “Production of synthetic rubbers” , M., 1967, p. 81).

The disadvantage of this method is the complex hardware design of the reactor unit and the inability to create large capacity plants due to the difficulties in organizing the circulation of coarse catalyst in the reactor-regenerator system.

Close in technical essence to the proposed one is a method for the dehydrogenation of n-butane to butylenes in a fluidized bed reactor-regenerator system with a finely dispersed aluminum-chromium catalyst (II Yukelson, “The Technology of Basic Organic Synthesis”, M., Chemistry Publishing House, 1968, pp. .180). According to this method, the feedstock - the butane fraction - enters in liquid form into the annular space of the shell-and-tube heat exchanger-evaporator for evaporation due to the heat of the contact dehydrogenation gas. Next, the feed vapor is fed into the tube space of the shell-and-tube heat exchanger-heater, where it is also heated by contact gas heat to a temperature of 275 ° C and then sent to a tube furnace, in which it is overheated to a temperature of 530-550 ° C. From the furnace, superheated vapors of n-butane at a pressure of 0.15 MPa are fed into the reactor with a fluidized bed of catalyst. Dehydrogenation is carried out at a temperature of 580 ° C. Dehydrogenation contact gas leaving the reactor It is used as a heat carrier for evaporation of liquid raw materials and heating of vapor of raw materials. In this case, the contact gas flows countercurrently to the feed vapors, first in the annular space of the heat exchanger-vapor heater, and then in the tube space of the heat exchanger-evaporator. Then, the cooled contact gas enters the scrubber of water washing and cooling the contact gas and is directed to the release of butylenes.

A disadvantage of the known method is the use of contact gas in the heat exchanger-evaporator as a heat carrier, contaminated with catalyst dust removed from the fluidized bed, and also prone to condensation under the operating conditions of the evaporator by high-boiling hydrocarbons generated during the dehydrogenation process. The rapid contamination of the heat transfer surface of the heat exchanger-evaporator with deposits of catalyst and resins determines the unreliability and inefficiency of the evaporation unit of the feedstock and, as a result, the instability of the installation in the known method. A redundant evaporation system for raw materials is required to enable frequent cleaning of heat exchangers without stopping production. In addition, the supply of contaminated contact gas also to the annular space of the heat exchanger-heater leads to a deterioration in heat transfer also in the heat exchanger-heater when it is practically impossible to clean the deposits of the annular space of the heat exchanger-heater.

The closest in technical essence and the achieved result is a method for producing butylenes by dehydrogenation of n-butane and a method for producing isoamylenes by dehydrogenation of isopentane in a fluidized bed of finely dispersed aluminum-chromium catalyst circulating in a reactor-regenerator system, including the preparation of feedstock by mixing fresh and recycle streams of the corresponding paraffinic hydrocarbons form, evaporation of feedstock in an evaporator heated by hot process water, heating obtained of vapors of raw materials in quenching coils located in the separation zone of the reactor due to the heat of the contact gas to a temperature of 130-160 ° C and their overheating in the coils of the furnace to a temperature of 500-550 ° C due to the heat of combustion of the gaseous fuel supplied to the furnace with the subsequent direction superheated feed vapors into the reactor at dehydrogenation carried out at a temperature of 530-590 ° C, a pressure in the reactor of 0.125 MPa and a volumetric feed rate of 120-180 hours ¹ , also consisting in the output of the coked catalyst to the regenerator for burning coke and restoring the activity of the catalyst at a temperature of 640-650 ° C and pressure of 0.117 MPa in the presence of air, including also cooling the contact gas in the recovery boiler by evaporating water condensate to produce secondary water vapor, as well as in a scrubber irrigated with water, compressing the cooled contact gas, condensation of the paraffin-olefin fraction and the separation of the obtained olefin hydrocarbons, as well as unreacted paraffin hydrocarbons with the direction of the latter for recycling to dehydrogenation (I. L. Kirpichnikov, V. V. Beresnev, L. M. Popov “Album of technological schemes of the main industries synthetic rubber "," Chemistry ", Leningrad, 1986, pp. 8-14, 57-74).

However, quenching coils in the known methods located in the separation zone of the reactor have a low heat transfer coefficient due to the relatively low velocity of the gas flow of contact gas in the zone of location of the coils. In this regard, the use of quench coils in this method for heating raw material vapor requires placing a large heat exchange surface in a limited volume of the reactor separation zone, which determines the design complexity of the coils, their unreliability in operation and limits the ability to heat the raw material vapor. The low temperature of the raw material vapor after quenching coils (130-160 ° C) requires increasing the furnace capacity for further overheating of the raw material vapor in the known method (up to 500-550 ° C), which increases operating costs and worsens the production environment when the furnace flue gases are discharged into the atmosphere .

The location of the quenching coils in the separation zone of the reactor above the fluidized bed of the catalyst under practical dehydrogenation processes often leads to the contact of the fluidized bed with the heat transfer surface of the coils. This occurs, for example, during pressure pulsations in a fluidized bed, inhomogeneous boiling and emission of aggregates of catalyst particles into the superlayer space, as well as during fluctuations of the fluidized bed level during process control and in other situations. Besides Moreover, in the separation zone of the reactor above the quenching coils, a fluidized bed of coked catalyst is accumulated containing predominantly small separated fractions with limited recirculation of this part of the catalyst through the regenerator. This is especially evident when the plants operate at increased loads of raw materials, requiring an increase in the amount of catalyst and, accordingly, the level of the fluidized bed in the reactor. All this leads to cooling by the coils of the upper part of the fluidized bed of the reactor, a heat deficit for the implementation of the endothermic dehydrogenation reaction and the need to fill this deficit by forcing the regenerator operating modes and catalyst circulation in the reactor-regenerator system. In this case, the stability of the installation is violated, dehydrogenation rates are reduced, and catalyst losses increase.

A disadvantage of the known methods is the use in the heat exchanger-evaporator as a heat carrier of hot water contaminated with catalyst dust and resins taken from the circulating circuit of the scrubber of water washing and cooling of the contact gas. The rapid contamination of the heat transfer surface of the heat exchanger-evaporator with deposits of catalyst and resins determines the unreliability and inefficiency of the evaporation unit of raw materials in the known method. In addition, the operation of the scrubber at the specified temperature of the circulating water (82-98 ° C) is associated with a large ablation of water vapor with contact gas and requires large expenses for the subsequent condensation of this water.

The disadvantages of the known method also include the high temperature of the contact gas in front of the cooling scrubber (temperature at the outlet of the recovery boiler) which reaches 300-400 ° C, which requires large expenses for cooling the water circulating in the scrubber cooling circuit. At the same time, relatively high temperatures of the contact gas after the scrubber (in front of the compressor) do not allow the dehydrogenation units to operate at high feedstock loads, increase the pressure in the reactor, and reduce the yield of olefins to decomposed paraffins.

These disadvantages of the known methods limit the power of the newly created reactors, and also make it impossible to increase production olefinic hydrocarbons in existing installations on existing equipment. The desire to maximize the production of olefin hydrocarbons by increasing the reactor load on raw materials leads to a decrease in the yield of olefin hydrocarbons on the passed and decomposed raw materials and to increased losses of the pulverized catalyst. For example, when using IM-2201 catalyst, catalyst losses reach 25-30 kg per ton of olefins.

Disclosure of invention

An object of the present invention is to increase the productivity of C3-C5 paraffin hydrocarbon dehydrogenation units, increase the olefin hydrocarbon yields and reduce catalyst consumption.

A method is proposed for producing olefin hydrocarbons by dehydrogenation of paraffin hydrocarbons in a fluidized bed of a dusty aluminum-chromium catalyst circulating in a reactor-regenerator system, including evaporation of a paraffin-containing raw material, consisting of a mixture of fresh and recycled paraffin hydrocarbon streams, heating the vapor obtained by contact gas heat and overheating of the contact gas 7 followed by a direction for dehydrogenation, which also includes cooling the contact gas of dehydrogenation in a waste heat boiler 15 with water vapor evaporation due to evaporation of the water condensate, as well as in a scrubber 17 irrigated with water, compression of the cooled contact gas, condensation and recovery of the obtained olefinic hydrocarbons and unreacted paraffin hydrocarbons.

The raw material vapor is heated to a temperature of 150-180 ° C in a heat exchanger 4 installed after the waste heat boiler 15, and the dehydrogenation is carried out at a volumetric feed rate of 120-200 hours ¹ , a temperature of 530-600 ° C and a linear velocity of the contact gas at the outlet of fluidized bed 0.5 m / s and above.

In this case, the water vapor obtained in the recovery boiler 15 is directed mainly to the evaporation of paraffin-containing raw materials and / or to the boilers of the columns for the separation of olefin and unreacted paraffin hydrocarbons In this case, a shell-and-tube heat exchanger of the vertical type with the supply of contact gas to the tube space of the heat exchanger is mainly used as a heat exchanger 4.

The temperature of the contact gas at the inlet to the waste heat boiler 15 is 520-590 ° C.

The main differences of the proposed method from the known are:

- the exclusion of the use for heating the vapor of the raw materials of the quenching coil 21 in the reactor 9 and, accordingly, the above-mentioned disadvantages of the dehydrogenation processes associated with its use;

- increasing the temperature of the contact gas at the inlet to the recovery boilers to 520-590 ° C, increasing due to this generation of water vapor in the recovery boilers, including obtaining higher-quality water vapor suitable for heat use in a wide range of consumers of the dehydrogenation unit and related industries;

the use of water vapor to vaporize paraffin-containing raw materials with the exception of using hot water from a scrubber 17;

- the use of contact gas for heating the vapor of the raw material of the heat exchanger on the pipeline 5 after the waste heat boiler 15 with an increase in the temperature of the vapor of the raw material to 150-180 ° C;

- reducing the power of the furnace 7 for overheating of the feed vapor at the inlet to the reactor 9 due to the higher temperature of the feed vapor at the inlet to the furnace 7.

A heat exchanger — a heat exchanger for heating raw material vapors in the proposed method for carrying out dehydrogenation processes has the following advantages over its analogue — coils 21 for heating raw material vapors in the prototype:

- the heat exchanger has several times lower hydraulic resistance along the vapor path of the feedstock;

- the heat exchanger provides several times greater heat transfer coefficient;

- the heat exchanger with a more compact arrangement of the heat exchange pipes significantly exceeds the coil 21 in design capabilities increase the heat transfer surface and practically does not create restrictions when increasing the capacity of dehydrogenation reactors, also having greater reliability compared to coils.

The temperature of the contact gas at the inlet to the recovery boiler 15 (520-590 ° C) in the proposed method with the exception of the quenching coil 21 in the reactor 9 is determined by the top temperature of the fluidized bed in the dehydrogenation reactor of paraffin hydrocarbons (530-600 ° C) taking into account heat losses along the path contact gas to the entrance to the recovery boiler 15 and provides the maximum possible thermal potential of the contact gas for its further use according to the proposed heat recovery scheme, taking into account the requirements of consumers received in the recovery boiler 15 secondary water vapor of the necessary parameters (evaporator of raw materials, boilers of the recovery columns for olefin and unreacted paraffin hydrocarbons), as well as taking into account the emerging opportunities for saving fuel gas and improving other indicators of the furnace 7 overheating of raw vapor.

The range of increase in the temperature of the vapor of the raw material after the heat exchanger-heater 4 and, accordingly, at the inlet to the furnace 7 (150-180 ° C) is limited by the conditions of economic feasibility in the conditions of a specific production: on the one hand, the generation of water vapor of the required quantity and parameters in the waste heat boiler 15, and on the other hand, the magnitude of energy costs during overheating of the raw material vapor in the furnace 7.

The elimination of the use of quenching coils 21 in the reactor 9 for heating the raw material vapor, as well as the replacement of the coolant in the raw material evaporator (replacing hot water from the scrubber 17 with high-quality water vapor obtained in the recovery boiler 15), increase the stability of the dehydrogenation unit, and open up opportunities for increasing power plants, increase dehydrogenation and reduce catalyst consumption.

The shell-and-tube heat exchanger-heater 4 of the vertical type when supplying contact gas to the pipe space provides higher heat transfer efficiency, is compact when installing the equipment of a dehydrogenation unit, and is convenient when cleaning pipes from contaminants. The proposed technical solution has advantages over the prototype in the specified range of changes in the main parameters of the processes of dehydrogenation of paraffin hydrocarbons (volumetric feed rate of the feed to the reactor, dehydrogenation temperature and linear velocity of the contact gas at the outlet of the fluidized bed).

The volumetric feed rate is determined as the ratio of the volumetric feed rate to the reactor under normal conditions (nm ³ / h) to the volume of catalyst loaded into the reactor (m).

The linear velocity of the contact gas at the outlet of the fluidized bed is defined as the ratio of the amount of contact gas under operating conditions (m ³ / h) to the cross-sectional area of the reactor (m ² ). The contact gas at the outlet of the fluidized bed contains, in addition to the contact dehydrogenation gas itself, also auxiliary flows, such as, for example, nitrogen supplied to the desorption zone of the reactor, gas flows supplied to transport the catalyst circulating in the reactor-regenerator system, etc. this cross-sectional area of the reactor is determined minus the cross-section of structural elements of the internal devices of the reactor.

The combination of these parameters is sufficient to determine the critical height of the fluidized bed in the reactor and the mode of its operation (S.M. Komarov, V.I. Korobov, A.L. Zeilingold, “Model of expansion of a fluidized bed”, “Theoretical principles of chemical technology”, Volume XVII 1983, Ns 6, p. 808). The height of the fluidized bed, as well as the totality of the claimed parameters, is the main criterion for assessing the boundary operating conditions of the reactor.

The technical result of the invention is to increase the productivity of plants dehydrogenation of paraffin hydrocarbons With ₃ - C5, an increase in the yield of olefinic hydrocarbons and a decrease in catalyst consumption.

Brief Description of the Drawings

Figure 1 shows a diagram of a plant for the dehydrogenation of paraffin hydrocarbons C3-C5, 'illustrating the invention. The feedstock, which is a mixture of fresh paraffin hydrocarbons and paraffin hydrocarbons, is recycled, sent in liquid form via pipeline 1 under a pressure of 500-900 kPa, to evaporator 2, where it is evaporated and heated to a temperature of 70-100 ° C (to avoid condensation in the path evaporator - heat exchanger-heater) with high-steam supplied water vapor (at a pressure of saturated steam of water vapor of 600-1300 kPa). Next, the pairs of raw materials are sent through the pipeline 3 for further heating into the annular space of the shell-and-tube heat exchanger-heater 4, where they are heated with contact gas supplied through the pipeline 5 to the tube space of the heat exchanger. Heated vapors of the feedstock are removed from the heat exchanger 4 at a temperature of 150-180 ° C and a pressure of 250-450 kPa and then sent through a pipe 6 to the coils of the furnace 7, where they are heated to a temperature of 500-550 ° C by the heat of flue gases from the combustion of gaseous fuel 7 supplied through pipeline 8, and then sent to the reactor 9 under a fluidized bed of aluminum-chromium catalyst for dehydrogenation, carried out at a volumetric feed rate of 120-200 hours ¹ and a temperature of 530-600 ° C. The catalyst circulates in the reactor-regenerator system with the coked catalyst withdrawing from the reactor to the regenerator (not shown in Fig. 1) through the stripping zone of hydrocarbons trapped by the circulating catalyst in the lower part of the fluidized bed of reactor 9. Desorption is carried out by nitrogen supplied to the lower part of the stripping zone by pipeline 12. In the regenerator, coke is burned and the activity of the catalyst is restored in the presence of air at a temperature of 640-660 ° C and a pressure of 0.117 MPa with the return of the regenerated catalyst and regenerator to the reactor 9. The circulation is performed by katalizatoroprovodu 10 of the reactor 9 into the regenerator via the air supplied to the pneumatic transport, and katalizatoroprovodu 11 from the regenerator to the reactor 9 via the raw material supplied to the pneumatic transport of vapors. The catalyst and the transport gas are discharged into the upper part of the fluidized bed of the reactor 9 below the level of the fluidized bed 20 using the conditional distribution device 19 shown. Figure 1 also shows the location above the fluidized bed of the quenching coil 21 in the prototype circuit design. The obtained dehydrogenation contact gas is purified from catalyst dust in cyclones located in the separation zone of the reactor 9 (not shown in FIG. 1), it is removed from the reactor through a pipeline 5 and at a temperature of 520-590 ° C is supplied as a heat transfer medium to a recovery boiler 15 fed by water condensate through a pipeline 16. The resulting water vapor through a steam collector 14 are sent via pipeline 13 directly or through the factory network to the evaporator 2 for evaporation of the feedstock. Excess steam is sent via line 18 to other consumers of the dehydrogenation unit, such as, for example, units for separating a fraction of paraffin and olefin hydrocarbons from hydrocarbon condensate by distillation and / or an absorption unit for non-condensed hydrocarbons with subsequent release of hydrocarbons by desorption from the absorbent at a temperature of 190-193 ° С in a column with a boiler heated by steam obtained in waste heat boilers with a pressure of 1.3 MPa and a temperature of 193 ° C (not shown in Fig. 1). Next, the contact gas is sent as a heat carrier to the shell-and-tube heat exchanger-heater 4. After passing through the recovery boiler 15 and the heat exchanger 4, the contact gas is sent to the scrubber 17 for water washing and cooling, after which the contact gas cooled to a temperature of 35-40 ° С is fed to the grocery the compressor and further to the unit for condensation and separation of the obtained olefinic hydrocarbons (not shown in Fig. 1). The recovered unreacted paraffin hydrocarbons are recycled to prepare the feedstock.

The best embodiment of the invention

The method is illustrated by the following examples.

Dehydrogenation is carried out using an aluminum-chromium microspherical catalyst AOK 73-24 with a mass fraction of hexavalent chromium of 0, 8-3, 5%, with an average particle diameter of 45-90 microns.

The diameter of the reactor and the regenerator in the examples is 4.6-5, 1 m

The surface of the quenching coils 21 in the separation zone of the reactor 9 when implementing the known method is 112-128 m

A typical technological scheme of the dehydrogenation unit according to the known method involves the use of two, arranged in series on the line of contact gas, heat recovery boilers with a heat transfer surface of 460-495 m ² .

When implementing the proposed method, the second in the direction of the contact gas waste heat boiler 15 is used as a heat exchanger — a heater for vapor of raw materials.

The natural gas supplied to the nozzles of the furnace 7 contains 82-90 wt.% Methane.

During the dehydrogenation processes, equipment is cleaned on the contact gas line (heat recovery boilers and heat exchanger-heater 4 vapors of raw materials) from catalyst dust deposits on heat transfer surfaces by known methods, for example, periodic, once every 7-10 days, purging of heat-exchanging pipes with gas suspension equilibrium catalyst from the pneumatic conveying system of the catalyst circulation.

Examples 1, 3, 5.

The dehydrogenation of n-butane to butylenes is carried out according to the known method in an installation containing a reactor 9 and a regenerator with a diameter of 4.6 m, a quenching coil 21 in a reactor 9 with a heat transfer surface of 112 m ² and two recovery boilers with a surface of 460 m each, installed in series on contact gas lines.

The fresh catalyst AOK 73-24 is loaded and further loaded into the reactor during the process with a hexavalent chromium content of 3.5 wt.%. Dehydrogenation is carried out at a volumetric feed rate of 200 hours ^{' 1} at a temperature of 600 ° C and a pressure in the reactor of 0.150 MPa (0.50 ati). The regeneration of the catalyst is carried out at a temperature of 650 ° C and a pressure of 0.145 MPa (0.45 MPa). The air flow rate into the regenerator is 19500 nm ³ / hour (example 1), 21500 nm ³ / hour (example 3) and 22500 nm ³ / hour (example 5). The consumption of feed vapor for pneumatic transport of the catalyst from the regenerator to the reactor is 910 kg / hour (example 1), 1000 kg / hour (example 3) and 1100 nm ³ / hour (example 5). The air flow rate for pneumatic transport of the catalyst from the reactor to the regenerator is 860 kg / hour (example 1), 900 kg / hour (example 3) and 970 nm ³ / hour (example 5). As a feedstock, a mixture of fresh and recycled butane fraction streams with a content of: n-butane - 92.5 wt.%, Isobutane - 3.5 wt.%, Butylenes - 3.0 wt.%, C ₃ - 0 hydrocarbons is used. 6 wt.%. The liquid feed evaporates and Heats up to 70 ° C. Then the raw material vapor is heated by the heat of the contact gas in the quenching coil 21 of the reactor 9, then overheated in the furnace 7 and fed to dehydrogenation. The contact gas is cooled in a quenching coil 21 of reactor 9, then in waste heat boilers and then in a scrubber 17 irrigated with water. The temperature of the contact gas after the scrubber 17 was 40-43 ° C.

Examples 2, 4, 6.

The dehydrogenation of n-butane to butylenes is carried out according to the proposed method in an installation comprising a reactor 9 and a regenerator with a diameter of 4.6 m, a waste heat boiler 15 with a surface of 460 m and a heat exchanger-heater 4 with a surface of 460 m installed in series on the contact gas line.

A fresh catalyst is loaded into the reactor and further loaded during the process with a hexavalent chromium content of 3.5 wt.%. Dehydrogenation is carried out at a volumetric feed rate of 200 hours ¹ at a temperature of 600 ° C and a pressure in the reactor of 0.150 MPa (0.50 ati). The regeneration of the catalyst is carried out at a temperature of 650 ° C and a pressure of 0.145 MPa (0.45 MPa). The air flow rate in the regenerator is 19500 nm ³ / hour (example 2), 20500 nm ³ / hour (example 4) and 21500 nm ³ / hour (example 6). The consumption of feed vapor for pneumatic transport of the catalyst from the regenerator to the reactor is 910 kg / hour (example 2) 960 kg / hour (example 4) and 1000 nm ³ / hour (example 6). The air flow rate for pneumatic transport of the catalyst from the reactor to the regenerator is 860 kg / hour (example 2), 880 kg / hour (example 4) and 920 nm ³ / hour (example 6). A mixture of fresh and recycle butane fraction streams containing: n-butane - 92.5 wt.%, Isobutane - 3.5 wt.%, Butylenes - 3.0 wt.%, C ₃ - 0 hydrocarbons, is used as feedstock. 6 wt.%. The liquid feed is vaporized and heated in a heat exchanger-heater (not shown in FIG. 1) to 70 ° C by supplying water vapor to the evaporator and heat exchanger-heater. Then the raw material vapor is heated by the heat of the contact gas in the heat exchanger-heater 4, then overheated in the furnace 7 and fed to dehydrogenation. The contact gas is cooled in a waste heat boiler 15, then in a heat exchanger-heater 4 and then in a scrubber 17 irrigated with water. The temperature of the contact gas after the scrubber 17 was 33-37 ° C.

Example 7 The dehydrogenation of isobutane to isobutylene is carried out according to a known method in an installation comprising a reactor 9 and a regenerator with a diameter of 4.6 m, a quenching coil 21 in a reactor 9 with a heat transfer surface of 112 m ² and two recovery boilers with a surface of 495 m each, installed on the contact gas line sequentially.

In the reactor 9 is loaded and then loaded during the process, a fresh catalyst AOK 73-24 with a content of hexavalent chromium of 1.5 wt.%. Dehydrogenation is carried out at a feed volumetric feed rate of 165 hours ¹ , at a temperature of 575 ° C and a pressure in the reactor of 0.150 MPa (0.50 ati). The regeneration of the catalyst is carried out at a temperature of 650 ° C and a pressure of 0.145 MPa (0.45 MPa). The air flow into the regenerator is 22500 nm ³ / hour. The consumption of feed vapor for pneumatic transport of the catalyst from the regenerator to the reactor is 1100 kg / h. The air consumption for pneumatic transport of the catalyst from the reactor to the regenerator is 970 kg / h. As the feedstock, a mixture of fresh and recycle streams of the isobutane fraction of the following composition is used, wt.%: Isobutane - 95.8, isobutylene - 1.8, n-butane and C ₃ - 2.4 hydrocarbons. Liquid raw materials evaporate and heat up to 70 ° С. Then the raw material vapor is heated by the heat of the contact gas in the quenching coil 21 of the reactor 9, then overheated in the furnace 7 and fed to dehydrogenation. The contact gas is cooled in a quenching coil 21 of reactor 9, then in waste heat boilers and then in a scrubber 17 irrigated with water. The temperature of the contact gas after the scrubber 17 was 42-45 ° C

Example 8

The dehydrogenation of isobutane to isobutylene is carried out according to the proposed method in an installation comprising a reactor 9 and a regenerator with a diameter of 4.6 m, a waste heat boiler 15 with a surface of 495 m and a heat exchanger-heater 4 with a surface of 495 m ² installed in series on the contact gas line.

In the reactor 9 is loaded and then loaded during the process, a fresh catalyst AOK 73-24 with a content of hexavalent chromium of 1.5 wt.%. Dehydrogenation is carried out at a volumetric feed rate of 165 hours ^-1 at a temperature of 575 ° C and a pressure in the reactor of 0.150 MPa (0.50 ati). The regeneration of the catalyst is carried out at a temperature of 650 ° C and a pressure of 0.145 MPa (0.45 MPa). The air flow into the regenerator is 21000 nm ³ / hour. Consumption of vapors of raw materials for catalyst pneumatic transport from the regenerator to the reactor - 1000 kg / h. The air consumption for pneumatic transport of the catalyst from the reactor to the regenerator is 910 kg / h. As the feedstock, a mixture of fresh and recycle streams of the isobutane fraction of the following composition is used, wt.%: Isobutane - 95.8, isobutylene - 1.8, n-butane and C ₃ - 2.4 hydrocarbons. The liquid feed is vaporized and heated in a heat exchanger-heater (not shown in FIG. 1) to 70 ° C by feeding steam to the evaporator 2 and the heat exchanger-heater. Then the raw material vapor is heated by contact gas heat in the heat exchanger-heater 4, then overheated in the furnace 7 and fed to dehydrogenation. The contact gas is cooled in a waste heat boiler 15, then in a heat exchanger-heater 4 and then in a scrubber 17 irrigated with water. The temperature of the contact gas after the scrubber was 35-40 ° C.

Example 9

The dehydrogenation of isopentane to isoamylenes is carried out according to a known method in an installation comprising a reactor 9 and a regenerator with a diameter of 5.1 m, a quenching coil 21 in the reactor 9 with a heat transfer surface of 128 m and two recovery boilers with a surface of 495 m ² each, installed on the contact gas line sequentially.

In the reactor 9 is loaded and then loaded during the process, a fresh catalyst AOK 73-24 with a hexavalent chromium content of 0.8 wt.%. Dehydrogenation is carried out at a volumetric feed rate of 120 hours ^{' 1} , at a temperature of 530 ° C and a pressure in the reactor of 0.1450 MPa (0.450 ati). The regeneration of the catalyst is carried out at a temperature of 660 ° C and a pressure of 0.14 MPa (0.4 ati). The air flow into the regenerator is 23500 nm ³ / hour. The consumption of raw material vapor for pneumatic transport of the catalyst from the regenerator to the reactor is 1150 kg / h. The air flow rate for pneumatic transport of the catalyst from the reactor to the regenerator is 1050 kg / h. As a feedstock, a mixture of fresh and recycle streams of the isopentane fraction of the following composition is used, wt.%: Isopentane - 97.0, isoamylenes - 1.8, hydrocarbons C ₄ - 1.0, hydrocarbons C ₆ and higher - 0.2. Liquid raw materials evaporate and heat up to 100 ° С. Then the raw material vapor is heated by the heat of the contact gas in the quenching coil 21 of the reactor 9, then overheated in the furnace 7 and fed to dehydrogenation. Contact gas cooled in quenching the coil 21 of the reactor 9, then in the waste heat boilers and then in the scrubber 17, irrigated with water, the temperature of the contact gas after the scrubber 17 to 43-47 ° C.

Example 10

The dehydrogenation of isopentane to isoamylenes is carried out according to the proposed method in an installation comprising a reactor 9 and a regenerator with a diameter of 5.1 m, a waste heat boiler 15 with a surface of 495 m ² and a heat exchanger-heater 4 with a surface of 495 m ² installed in series on the contact gas line.

In the reactor 9 is loaded and then loaded during the process, a fresh catalyst AOK 73-24 with a hexavalent chromium content of 0.8 wt.%. Dehydrogenation is carried out at a volumetric feed rate of 120 hours ¹ at a temperature of 530 ° C and a pressure in the reactor of 0.145 MPa (0.45 MPa). The regeneration of the catalyst is carried out at a temperature of 660 ° C and a pressure of 0.14 MPa (0.4 ati). The air flow into the regenerator is 22000 nm ³ / hour. The consumption of feed vapor for pneumatic transport of the catalyst from the regenerator to the reactor is 1050 kg / h. Air consumption for pneumatic transport of the catalyst from the reactor to the regenerator is 980 kg / h. As a feedstock, a mixture of fresh and recycle streams of the isopentane fraction of the following composition is used, wt.%: Isopentane - 97.0, isoamylenes - 1.8, hydrocarbons C ₄ - 1.0, hydrocarbons C ₆ and higher - 0.2. The liquid feed is vaporized and heated in a heat exchanger-heater (not shown in FIG. 1) to 100 ° C by feeding steam to the evaporator 2 and the heat exchanger-heater. Then the raw material vapor is heated by the heat of the contact gas in the heat exchanger-heater 4, then overheated in the furnace 7 and fed to dehydrogenation. The contact gas is cooled in the waste heat boiler 15, then in the heat exchanger-heater 4 and then in the scrubber 17, irrigated with water, the temperature of the contact gas after the scrubber 17 to 37-40 ° C.

The main conditions for the implementation of the processes and the achieved dehydrogenation rates are presented in tables 1 and 2.

From the above conditions for the implementation of the processes and tables 1, 2 it can be seen that the air flow to the regenerator, the flow of vapors of raw materials and air for the circulation of the catalyst in the reactor-regenerator system and, as a consequence, the costs of other auxiliary flows in the dehydrogenation unit (for figure 1) - the flow rate of natural gas to the regenerator in the combustion zone — heating the catalyst, natural gas to the catalyst recovery zone, the nitrogen flow to the desorption zones of the reactor and the regenerator, as well as the dehydrogenation rates at relatively low loads of a typical dehydrogenation unit for raw materials (22 t / h) and, correspondingly, at linear velocities of the contact gas at the outlet of the fluidized bed of the catalyst of less than 0.5 m / s, both the prototype and the invention are almost identical. At the same time, with an increase in the load of the dehydrogenation unit for raw materials (up to 27-44 t / h) and, correspondingly, at linear velocities of the contact gas at the exit from the fluidized bed of the catalyst of more than 0.5 m / s, the indicated flow rates during dehydrogenation processes by the invention are reduced, the yields of olefins on the passed and decomposed paraffin hydrocarbons are increased and the specific consumption of the catalyst is reduced compared with the prototype. The examples show that typical dehydrogenation plants based on the technology of the prototype using two consecutive recovery boilers (waste-heat boiler 1 and waste-heat boiler 2, see tables 1 and 2), are easily reconstructed to switch to the proposed technology by using the second along the contact gas of the recovery boiler (recovery boiler 2) as a heat exchanger 4 for heating the vapor of the feed (gas-gas heat exchanger, see tables 1 and 2). As can be seen from the tables, in the implementation of the known method in the second along the contact gas recovery boiler 15, a relatively small amount of water vapor is generated, which confirms the advisability of using this recovery boiler as a heat exchanger-heater 4 vapor of raw materials when transferring a typical installation to the proposed method of implementation dehydrogenation processes. It is shown that on the basis of the available equipment of a typical installation, it is possible to switch to the use of the proposed technology with a significant increase in plant capacity and an improvement in dehydrogenation. At the same time, it is also possible to increase the production of water vapor in the waste heat boiler and reduce the consumption of fuel gas and the amount of waste flue gases from the furnace for overheating of the feed vapor. A significant decrease in the temperature of the contact gas at the inlet to the scrubber 17 in comparison with the prototype allows to stabilize the work a scrubber 17 in a narrow temperature range with a decrease in the temperature of the contact gas after the scrubber 17 to an economically feasible temperature from the point of view of minimizing energy costs for cooling the contact gas.

The exclusion of quenching coils 21 of the reactor 9 does not lead to undesirable conversions of the contact gas in the separation zone of the reactor in the conditions of the proposed method. The formation of thermopolymers and clogging of the annular space in the heat exchanger-heater 4 vapor of raw materials is not observed.

Thus, the use of the proposed method for producing olefinic hydrocarbons allows to stabilize the operation of dehydrogenation plants of paraffin hydrocarbons, increase the productivity of plants, increase the yields of olefin hydrocarbons and reduce the consumption of catalyst in comparison with the prototype.

Industrial applicability

The method can be used for the production of basic monomers of synthetic rubbers, esters, as well as other organic products.

Claims

CLAIM

1. A method of producing olefinic hydrocarbons by dehydrogenation of paraffin hydrocarbons in a fluidized bed of a dusty alumino-chromium catalyst circulating in a reactor-regenerator system, comprising vaporizing a paraffin-containing raw material consisting of a mixture of fresh and recycled paraffin hydrocarbon streams, heating the resulting vapors by contact gas heat and overheating in furnace (7) with a subsequent direction for dehydrogenation, which also includes cooling the contact gas dehydrogenation in a waste heat boiler (15) to obtain secondary water vapor due to the evaporation of water condensate, as well as in a scrubber (17), irrigated with water, compression of the cooled contact gas, condensation and evolution of the obtained olefinic hydrocarbons and unreacted paraffin hydrocarbons, characterized in that the feed vapor is heated to a temperature of 150-180 ° C in a heat exchanger (4) installed after the recovery boiler (15), while dehydrogenation is carried out at a volumetric feed rate of 120-200 hours ¹ , a temperature of 530-600 ° C and a linear velocity of the contact gas to the exit e from a fluidized bed of 0.5 m / s and above.

2. The method according to claim 1, characterized in that the water vapor obtained in the recovery boiler (15) is directed to the evaporation of the paraffin-containing feedstock and / or to the boilers of the recovery columns for olefin and unreacted paraffin hydrocarbons.

3. The method according to any one of claims 1 to 2, characterized in that a shell-and-tube heat exchanger of a vertical type is used as the heat exchanger (4), while the contact gas is fed into the tube space of the heat exchanger.

4. The method according to claim 1, characterized in that the temperature of the contact gas at the inlet to the waste heat boiler (15) is 520-590 ° C.