KR101562691B1 - Hydrocarbon dehydrogenation process - Google Patents

Abstract

The present invention relates to a catalytic dehydrogenation process of C3 + feed hydrocarbons. The hydrocarbon feedstock is initially split and the first portion of the feedstock is introduced into the first dehydrogenation reaction zone where it is operated without oxidative reheating and the resulting effluent is subjected to a second dehydrogenation Is introduced into the reaction zone. The resulting effluent from the second dehydrogenation reaction zone is introduced into the third dehydrogenation reaction zone, which is operated with oxidative reheat with the second portion of the feedstock.

Description

Hydrocarbon dehydrogenation process {HYDROCARBON DEHYDROGENATION PROCESS}

The technical field to which the present invention relates is catalytic dehydrogenation of hydrocarbons. More specifically, the present invention is a process for dehydrogenating hydrocarbons including paraffins and alkyl aromatics.

Improvements in existing contact dehydrogenation plants that do not have oxidative reheat capability to increase capacity are limited by the reactor internal velocity and the maximum temperature of the reactor transfer line. Due to these limitations in current installations, the maximum possible capacity expansion by simply increasing throughput in current reactors is 35%.

Dehydrogenation of hydrocarbons is well known, and therefore both acrylic and aromatic hydrocarbons are converted to the corresponding less saturated products. As a practical example, dehydrogenation is typically carried out to produce styrene from ethylbenzene. US 3,515,766 and US 3,409,689 disclose steam dehydrogenation processes for alkyl aromatics, including ethylbenzene. These references describe a mixture of an extra amount of superheated steam and reactant between successive layers of a mixture of superheated steam and feed hydrocarbons and a dehydrogenation catalyst to reheat the reactants.

The prior art also teaches that oxygen is passed through the dehydrogenation zone, which is intended to release heat and consume oxygen by reacting free hydrogen and oxygen during the dehydrogenation reaction. Processes known for using this technique utilize a hydrogen oxidation catalyst in an attempt to selectively oxidize hydrogen rather than the feed or product hydrocarbons present in the dehydrogenation zone as well.

Summary of the Invention

The present invention provides a means for increasing the capacity of current two-reactor dehydrogenation processes operated without oxidative reheating. Thus, the hydrocarbon feedstock is initially split, the first portion of the hydrocarbon feedstock is introduced into the first dehydrogenation reaction zone, which is operated without oxidative reheating, and the resulting effluent is reheated sequentially, Is introduced into the second dehydrogenation reaction zone which is operated without reheating. The resulting effluent from the second dehydrogenation reaction zone is introduced with a second portion of the hydrocarbon feedstock into the third dehydrogenation reaction zone, which is operated by oxidative reheat.

The portion of the feed that bypasses the first two dehydrogenation reaction zones is basically the feed volume that achieves the desired capacity increase. For example, when 33% of the total hydrocarbon feed is bypassed, a 50% increase in total plant capacity results if two reaction zones start and reach three reaction zones. When performing the entire process in this way, there is no change in the steam flow rate, the temperature for steam soaking and the combined feed for the current first dehydrogenation reaction zone remain unchanged from the original design, Important compartments, such as steam superheaters and two existing reactors, are not directly affected by capacity expansion. A 50-60% capacity increase is feasible using the process of the present invention.

Brief Description of Drawings

The figure is a simplified process flow diagram of a preferred embodiment of the present invention. The drawings are intended to illustrate the invention in a schematic but not to limit the invention.

DETAILED DESCRIPTION OF THE INVENTION

The dehydrogenation process of aromatic hydrocarbons exists in a wide range of conventional applications. As a practical example, a large amount of styrene is produced by dehydrogenation of ethylbenzene. The resulting styrene may be polymerized on its own, or the styrene may be copolymerized with butadiene, isoprene, acrylonitrile, and the like. Other hydrocarbons which can be dehydrogenated in the same way in large amounts include diethylbenzene, ethyltoluene, propylbenzene, and isopropylbenzene. The subject process can also be applied to the dehydrogenation of other types of hydrocarbons, including relatively pure or mixed streams of C2-C16 paraffins. Therefore, the process can also be applied to the dehydrogenation of propane, butane, hexane or nonane. However, since the majority of current conventional dehydrogenation processes are used for the dehydrogenation of ethylbenzene, the description of the subject invention described later will be presented mainly in terms of dehydrogenation of ethylbenzene. This is not intended to exclude those alkylaromatic and acyclic hydrocarbons described above from the scope of the subject invention or those having different ring structures, including bicyclic compounds.

The dehydrogenation reaction is highly endothermic. Thus, the process of passing the reactants through the dehydrogenation catalyst bed results in a reduction of the reactant temperature. The endothermicity of the reaction is such that the temperature reduction removes the reactants from the predetermined temperature range. The reactants are actually cooled to such an extent that the desired reaction no longer proceeds at a normally achievable rate. Therefore, the per pass conversion desired or normally required per pass can not be achieved simply by contacting the reactants with a single layer of the dehydrogenation catalyst. For this reason, it has become a standard routine practice to perform interstage reheating in several ways. In the interheat reheating, the reaction effluent of the first catalyst bed is heated to the desired inlet temperature of the downstream second catalyst bed. This reheating may be performed by direct heat exchange, such as by mixing hot steam into the reactant stream from the first catalyst bed of the hot steam.

A preferred intergrowth reheat method involves the use of indirect heat exchange. In this way, the effluent from the dehydrogenation zone passes through a heat exchanger where it is heated, and then the reactants are introduced into the subsequent dehydrogenation zone. The hot fluid used in this indirect heat exchange process may be hot steam, a combustion gas, a hot process stream or other readily available hot fluid.

According to the present invention, a first portion of the dehydrogenatable hydrocarbon feedstock is heated and introduced into a first dehydrogenation reaction zone, preferably operated with steam, without oxidation reheating, to produce an effluent stream, The stream is reheated and introduced into the second dehydrogenation reaction zone. The resulting effluent from the second dehydrogenation reaction zone contains the hydrogen produced during the dehydrogenation and then the hydrogen is contact oxidized to generate heat and reheat the reactants prior to introduction of the reactants into the downstream dehydrogenation catalyst Available. The resultant effluent from the second dehydrogenation reaction zone, the second portion of the dehydrogenatable hydrocarbon feedstock, oxygen and optionally steam are reacted in a third dehydrogenation reaction zone operated by oxidative reheating.

The second portion of the dehydrogenatable hydrocarbon feedstock, oxygen and, if present, steam is subjected to a third dehydrogenation process to ensure a homogeneous feed to the oxidative reheat catalyst to achieve the desired reaction temperature before contacting the dehydrogenation catalyst It is preferable that they are mixed before entering the reaction zone. This uniformity also eliminates the probability of having an explosive concentration of hydrocarbons.

The motivation for using oxidative reheating is the recognition that the combustion of hydrogen generated in the dehydrogenation reaction zone performs two functions that are beneficial to the dehydrogenation process. First, the consumption of hydrogen is beneficial in shifting the equilibrium of the dehydrogenation reaction, which predominates the increased amount of dehydrogenation. Second, selective combustion of hydrogen releases sufficient heat to reheat the reactants under the desired dehydrogenation conditions.

The oxidation is preferably achieved in the presence of a catalyst which selectively promotes the oxidation of hydrogen as compared to destructive combustion or oxidation of more efficient feed and product hydrocarbons. The selective combustion method of inter-reheat provides a more economical dehydrogenation process.

Despite the progress achieved in the art of catalysis and hydrocarbon conversion, the ultimate conversion rate that can be achieved during a single pass through the dehydrogenation zone is limited to less than the total conversion rate. That is, it is impossible to achieve 100% conversion from the feed hydrocarbons to the corresponding product dehydrogenated hydrocarbons. The fundamental limitation in the degree of conversion that can be achieved in any dehydrogenation process is the equilibrium concentration of the various reactants at the temperature used. Therefore, the effluent stream in the catalytic dehydrogenation zone comprises a mixture of feed hydrocarbons, dehydrogenated hydrocarbon products, and hydrogen. In general, it is necessary to separate and recover the dehydrogenated hydrocarbon product and recycle the unconverted feed hydrocarbons. The higher the conversion achieved in the dehydrogenation zone is, the smaller the amount of recycled unconverted material is. The separation of product and unconverted hydrocarbons requires extensive capital equipment and consumes large amounts of equipment in the form of heat and electric power. It is therefore desirable to reduce the amount of material that must be separated and recycled by increasing the conversion achieved per pass in the dehydrogenation zone. The high conversion per pass rate also allows for smaller reaction zones that can be used in the process in connection with the reduction in reactor cost, the catalyst and equipment cost of operating the reaction zone. For this reason, it is desirable to achieve an increased proportion of the overall conversion rate during the passage of the dehydrogenated zone feed stream through the multi-layer dehydrogenation zone.

In the oxidative reheating step, it is preferred that the oxygen-containing gas stream is mixed with the effluent of the preceding dehydrogenation reaction zone and the resulting mixture is introduced into the layer of the selective oxidation catalyst together with a portion of the dehydrogenatable hydrocarbon feedstock do. In order to achieve an optimum level of performance and stability in such a process, it is necessary to finely control the rate at which oxygen is introduced into the process in such a manner.

An insufficient amount of oxygen results in less consumption than the consumption of the desired hydrogen and, more important, less reheating than reheating the desired reactant stream. The result is a reduction in the degree of dehydrogenation achieved during passage over the entire reaction zone. In general, it is not desirable to inject an excess of oxygen over any portion of the dehydrogenation zone that is required to perform the desired degree of hydrogen combustion.

Injecting excess oxygen into the dehydrogenation zone may also have a detrimental effect on long-term operation of the process. In practice, for example, oxygen generally functions to deactivate or partially neutralize some commonly used dehydrogenation catalysts. Therefore, it is not desirable to have residual oxidation oxygen from the oxidation catalyst layer and the dehydrogenation catalyst contacting it on it. Operation of the dehydrogenation zone in a manner that does not result in total consumption of oxygen is also undesirable because of the apparent explosive nature of the oxygen-hydrocarbon mixture. However, the explosive nature of such a mixture is not sufficient to avoid the presence of a mixture which is basically within the explosion range, such as using a diluent and a sufficient amount of solid material to act as an intentional low oxygen addition rate and explosion control means It can be invalidated by properly operating it. Finally, the presence of oxygen is generally undesirable in hydrocarbon-containing vessels because oxygen can react with hydrocarbons to form various undesired oxyacid compounds.

The structure of the dehydrogenation reaction zone can be varied by changing the type of catalyst layer used. Actually, for example, there is a radial flow through the annular catalyst bed as well as a vertical flow through the cylindrical catalyst bed. It should be noted that, with radial flow, the layers of the dehydrogenation catalyst and the layers of the oxidation catalyst may be located in the same highly concentric circles within the vessel or vessel (s). Either an oxidation catalyst or a dehydrogenation catalyst may be located in the outer layer of this alignment. The gas flow then passes through the annular gas collecting and distributing pore volume located through the cylinder center pipe region located in the middle of the radial flow catalyst bed and between the outer surface of the catalyst bed and the inner wall of the vessel. Variations are also possible in the number of catalyst layers that can be used. Suitable systems for catalyst placement are described in US 3,498,755; US 3,515,763; And US 3,751,232.

The dehydrogenation catalyst generally comprises at least one metal component selected from Groups VI and VIII of the Periodic Table. One typical catalyst for the dehydrogenation of alkylaromatics comprises 85% by weight iron (III) oxide, 2% by weight chromia, 12% by weight potassium hydroxide and 1% by weight sodium hydroxide. A second dehydrogenation catalyst commonly used comprises 87 to 90% by weight of iron (III) oxide, 2 to 3% by weight of chromium oxide and 8 to 10% by weight of potassium oxide. A typical third catalyst comprises 90 wt% iron oxide, 4 wt% chromia, and 6 wt% potassium carbonate. Methods for preparing suitable catalysts are well known in the art. This is evidenced by the teachings of US 3,387,053, which discloses that the patent contains at least 35% by weight of iron oxide as catalyst, from 1 to 8% by weight of zinc or copper oxide, from 0.5 to 50% by weight of alkali promoter, The preparation of catalyst complexes of from 1 to 5% by weight of chromium (III) is described. US 4,467,046 also describes a catalyst for the dehydrogenation of ethylene benzene in the presence of steam. This catalyst contains 15 to 30% by weight of potassium oxide, 2 to 8% by weight of cerium oxide, 1.5 to 6% by weight of molybdenum oxide, 1 to 4% by weight of calcium carbonate and the balance of iron oxide.

In general, the dehydrogenation conditions include a temperature of 500 to 750 占 폚, preferably 565 to 675 占 폚. The temperature required for efficient operation of any particular dehydrogenation process depends on the activity of the feed hydrocarbons and the catalyst used. The pressure maintained in the dehydrogenation zone can range from 100 to 750 mmHg, and the pressure range is preferably from 250 to 700 mmHg. The operating pressure within the dehydrogenation zone is measured at the inlet, middle section and outlet of the zone to provide an average pressure. Feed stream is in the range of 0.1 to 2.0 hr ^-1, and preferably is charged to the dehydrogenation zone at a space velocity per hour of the liquid of 0.1 to 1.0 hr ^-1 based on the total liquid hydrocarbon charge at 15.6 ℃.

The hydrocarbon feed to be dehydrogenated is preferably mixed with the superheated steam to prevent the temperature lowering effect of the endothermic dehydrogenation reaction. The presence of steam is also described as benefiting the stability of dehydrogenation by preventing accumulation of carbon deposits. Preferably, the steam is mixed with the other components of the feed stream at a rate of 0.5 to 1.5 pounds of steam per pound of feed hydrocarbons. Other amounts of steam may be added after one or more subsequent dehydrogenation catalyst layers if desired. However, the dehydrogenated zone effluent stream contains less than 3 pounds of steam per pound of product hydrocarbon, preferably less than 2 pounds of steam per pound of product hydrocarbon.

 The vapor phase effluent stream derived from the final dehydrogenation zone may be heat exchanged for a stream of steam, a reactant stream of such process or other process, or it may be used as a heat source for fractional distillation. Typically, the effluent stream is often passed through several heat exchangers, thereby heating a number of different streams and cooling the effluent stream. This heat exchange is performed under the condition of the set limit. Preferably, the cooling is sufficient to condense C6 + hydrocarbons, i. E. At least 95 mol% of hydrocarbons having at least 6 carbon atoms per molecule and at least 95 mol% of water vapor in the dehydrogenated zone effluent stream. Thereby, all of the dehydrogenated hydrocarbon products, such as styrene, mostly water, and other readily condensable compounds present in the effluent stream, are basically converted to liquids. This produces a mixed phase stream that is introduced into the phase separation vessel. This procedure allows for easy crude separation by sloping hydrocarbons from water and hydrogen present in the effluent stream. The dehydrogenated hydrocarbon product present in the dehydrogenated zone effluent stream is part of the hydrocarbon stream that is recovered from the separation vessel and is transferred to the appropriate separation facility. The dehydrogenated hydrocarbon product is preferably recovered from the hydrocarbon stream using one of several fractionation systems known in the art. It is preferred that this fractional distillation yields a relatively pure stream of unconverted hydrocarbon feed such as ethylene benzene that can be regenerated for improved economy. Additional hydrocarbon streams, including by-products of the dehydrogenation reaction, can also be obtained during product preparation. For example, in the production of styrene from ethylene benzene, benzene and toluene may be recovered and partially recycled as taught in US 3,409,689 and GB 1,238,602, or may be totally rejected from the process. If desired, methods other than fractional distillation may be used to recover the dehydrogenated hydrocarbon product. As a practical example, US 3,784,620 teaches the separation of styrene and ethylbenzene through the use of polyimide permeable membranes such as nylon-6 and nylon 6,10. US 3,513,213 teaches a separation process using liquid-liquid extraction in which anhydrous silver fluoroborate is used as a solvent. A similar separation process using copper (II) fluoroborate and copper (II) fluorophosphate is described in US 3,517,079; US 3,517,080; And US 3,517,081.

 The oxygen feed stream to the process can be air, but is preferably a gas that has more oxygen content than air. The oxygen feed stream may have a nitrogen content of less than 10 mol%, and the use of substantially pure oxygen is highly desirable if it is economically viable. The preferred oxygen concentration in the oxygen feed stream is primarily an economic problem and can be determined by combining the advantages of having pure oxygen with the cost of obtaining that oxygen. A fundamental disadvantage of the presence of nitrogen is the dilution of the hydrogen-containing gas stream removed from the product separation vessel and the increase in absolute pressure maintained in the pressure drop and dehydrogenation zone across the catalyst bed as nitrogen is passed through the dehydrogenation zone. On the other hand, the presence of nitrogen advantageously affects the level of equilibrium conversion favorably by acting as a diluent.

The oxidation catalyst used in the oxidation reheat or oxidation zone to promote hydrogen oxidation may be any conventionally suitable catalyst. The oxidation catalyst has a composition different from that of the dehydrogenation catalyst. Preferably, the oxidation catalyst has a high selectivity for oxidation of hydrogen, with only a small amount of feed or product hydrocarbons being oxidized. Preferred oxidation catalysts include IUPAC 7,8 or 9 group noble metals and one or more other metal or metal cations, both of which are present in minor amounts on the refractory solid support. Preferred noble metals are platinum and palladium, but the use of ruthenium, rhodium, osmium and iridium is also contemplated. In an embodiment, the noble metal is present in an amount ranging from 0.01 to 5.0% by weight of the final catalyst. The metal or metal cation is selected from the group IUPAC 1 or 2 and is preferably present in an amount ranging from 0.01 to 20 wt% of the final catalyst. The metal or metal cation may be selected from the group consisting of lithium, potassium, rubidium and cesium. In an embodiment, the metal or metal cation is lithium or potassium. Another optional component of the oxidation catalyst may be selected from the group of IUPAC 14.

In a preferred embodiment, the refractory solid support of the oxidation catalyst has a surface area of from 1 to 300 m ² / g; An apparent bulk density of 0.2 to 1.5 g / cc; Alumina having an average pore size of 20 ANGSTROM or more. The metal-containing component is preferably impregnated into the solid particles of the solid support by dipping in an aqueous solution followed by drying and calcining at a temperature in the range of from 500 to 1200 < 0 > C in air. The support may be in the form of spheres, pellets or extrudates. The total amount of oxidation catalyst present in the dehydrogenation zone is preferably less than 30 wt% of the total amount of dehydrogenation catalyst, more preferably 5 to 15 wt% of the total amount of such dehydrogenation catalyst.

The conditions used during contacting the reactant stream with the oxidation catalyst layer are set to a large extent by the dehydrogenation conditions described above. The outlet temperature of the oxidation catalyst is the preferred inlet of the downstream dehydrogenation catalyst layer. The temperature rise across the oxidation catalyst is preferably controlled by the amount of hydrogen conversion across the oxidation catalyst. It is preferred that the space velocity per liquid hour based on liquid hydrocarbon filling at 15.6 占 폚 is 2 to 20 hr < ^-1 >.

Brief Description of Drawings

In the drawings, the process of the present invention is illustrated by a simplified schematic flow diagram in which details such as pumps, devices, heat exchange and heat recovery circuits, compressors, and the like are omitted where the details are not essential to an understanding of the relevant techniques. The use of such a variety of equipment is obviously within the scope of those skilled in the art.

Referring again to the drawings, a hydrocarbon feed stream comprising C3 + feed hydrocarbons, i.e., hydrocarbons having three or more carbon atoms per molecule, is introduced into the process via line 1 and is bifurcated into a first portion and a second portion, Two parts are provided. A first portion of the feed stream is conveyed through line 2 and combined with the steam provided by line 3, and the resulting mixture is introduced into dehydrogenation reaction zone 5. The dehydrogenated zone 5 is operated without oxidative reheating and the resulting effluent is conveyed through line 6 and heated via indirect heat exchange (not shown), and the dehydrogenation reaction zone 7). The dehydrogenated zone 7 is operated without oxidative reheating and the resulting effluent is combined with the second portion of the feed stream being transported through line 8 and through lines 13 and 9, . The resulting effluent from the dehydrogenation zone 7 carried through line 8 is also combined with the oxygen and steam mixture provided through lines 14 and 9. The resulting mixture is conveyed through line 10 and introduced into dehydrogenation reaction zone 11, which is carried out by oxidative reheating. The dehydrogenation reaction zone (11) contains an oxidation zone (16) and a dehydrogenation zone (17). The resulting effluent from dehydrogenation reaction zone (11) contains product dehydrogenated hydrocarbons carried and recovered via line (12).

Claims (10)

As a catalytic dehydrogenation method of C3 + feed hydrocarbons,
(a) passing a first portion of a feed stream comprising C3 + feed hydrocarbons to a first dehydrogenation catalyst bed under dehydrogenation conditions in a first dehydrogenation zone to produce a feed stream comprising hydrogen, C3 + feed hydrocarbons and C3 + product hydrocarbons Producing a first dehydrogenated zone effluent stream,
(b) at least a portion of the first dehydrogenated effluent stream is heated and passed through a second dehydrogenation catalyst bed under dehydrogenation conditions in a second dehydrogenation zone to produce a feed stream comprising hydrogen, a C3 + feed hydrocarbons and a C3 + 2 dehydrogenated zone effluent stream,
(c) passing at least a portion of the second dehydrogenated effluent stream, a second portion of the feed stream comprising C3 + feed hydrocarbons, and oxygen into an individual layer of a selective hydrogen oxidation catalyst under oxidizing conditions in the oxidizing region, Generating an effluent,
(d) passing at least a portion of the oxidation zone effluent through a third dehydrogenation catalyst bed under dehydrogenation conditions in a third dehydrogenation zone to produce a third dehydrogenation zone effluent stream comprising product hydrocarbons; and
(e) recovering the product hydrocarbons
≪ / RTI > The method of claim 1, further comprising mixing steam with a first portion of the feed stream and a first portion of the feed stream prior to passing steam through the first dehydrogenation catalyst bed. 3. The process of claim 2, wherein the amount of steam ranges from 0.5 to 1.5 pounds per pound of C3 + feed hydrocarbons in the first portion of the feed stream. 3. A process according to claim 1 or 2, wherein steam is introduced into the second dehydrogenated zone effluent stream, a portion of the second dehydrogenated zone effluent stream, a second portion of the feed stream and oxygen, Oxygen and steam are passed through the individual layers of the selective hydrogen oxidation catalyst. The process of any one of claims 1 to 3 wherein the C3 + feed hydrocarbons are alkylaromatic hydrocarbons. 3. The process according to claim 1 or 2, wherein the C3 + feed hydrocarbons are ethyl benzene. 3. The process according to claim 1 or 2, wherein the C3 + feed hydrocarbons are propane. 3. The process according to claim 1 or 2, wherein the C3 + feed hydrocarbons are butane. The process according to any one of claims 1 to 5, wherein the dehydrogenation conditions comprise a temperature of 500 ° C to 750 ° C and a pressure of 100 to 750 mmHg. 10. The process of claim 9, wherein the dehydrogenation conditions comprise a pressure of 250 to 700 mmHg.

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