WO2010101571A1

WO2010101571A1 - Hydrocarbon dehydrogenation process

Info

Publication number: WO2010101571A1
Application number: PCT/US2009/036129
Authority: WO
Inventors: Sanjeev Ram; Ajaykumar C. Gami; Guy B. Woodle
Original assignee: Uop Llc; Lummus Technology, Inc.
Priority date: 2009-03-05
Filing date: 2009-03-05
Publication date: 2010-09-10
Also published as: CA2753127C; RU2505516C2; RU2011140326A; CN102341484A; JP5638013B2; KR20110133473A; JP2012519685A; CN102341484B; CA2753127A1; MX2011009207A; KR101562691B1; BRPI0924262A2

Abstract

A process for the catalytic dehydrogenation of a C3+ feed hydrocarbon. The hydrocarbon feedstock is initially divided and a first portion of the feedstock is introduced into a first dehydrogenation reaction zone operated without oxidative reheat and the resulting effluent is introduced into a second dehydrogenation reaction zone operated without oxidative reheat. The resulting effluent from the second dehydrogenation reaction zone is introduced along with a second portion of the feedstock into a third dehydrogenation reaction zone operated with oxidative reheat.

Description

HYDROCARBON DEHYDROGENATION PROCESS

FIELD OF THE INVENTION

[0001] The field of art to which this invention pertains is the catalytic dehydrogenation of hydrocarbons. More specifically, the invention is a method for dehydrogenating hydrocarbons, including, paraffins and alkylaromatics.

BACKGROUND OF THE INVENTION

[0002] The revamp of an existing catalytic dehydrogenation facility having no oxidative reheating capability in order to increase capacity is limited by the reactor internal velocities and maximum temperatures of the reactor transfer lines. Because of these limitations on an existing facility, the maximum possible capacity expansion by simply increasing the throughput to the existing reactors is 35 percent.

[0003] The dehydrogenation of hydrocarbons is well known, with both acyclic and aromatic hydrocarbons being thereby converted to the corresponding less saturated products. For instance, dehydrogenation is performed commercially for the production of styrene from ethylbenzene. US 3,515,766 and US 3,409,689 disclose catalytic steam dehydrogenation processes for alkylaromatics including ethylbenzene. These references describe the admixture of superheated steam and the feed hydrocarbon and the admixture of additional amounts of superheated steam with the reactants between sequential beds of dehydrogenation catalyst to reheat the reactants. [0004] The prior art also teaches to pass oxygen into a dehydrogenation zone for the purpose of reacting the oxygen with hydrogen released during the dehydrogenation reaction to thereby liberate heat and to consume hydrogen. The processes known to employ this technique utilize a hydrogen oxidation catalyst in an attempt to selectively oxidize the hydrogen rather than feed or product hydrocarbons also present in the dehydrogenation zone.

SUMMARY OF THE INVENTION

[0005] The present invention provides a means for increasing the capacity of an existing two reactor dehydrogenation process which is operated without oxidative reheat. Accordingly, the hydrocarbon feedstock is initially divided and a first portion of the hydrocarbon feedstock is introduced into a first dehydrogenation reaction zone which is operated without oxidative reheat and the resulting effluent is subsequently reheated and introduced into a second dehydrogenation reaction zone which is also operated without oxidative reheat. The resulting effluent from the second dehydrogenation reaction zone is introduced, along with a second portion of the hydrocarbon feedstock, into a third dehydrogenation reaction zone which is operated with oxidative reheat.

[0006] The portion of the feed which bypasses the first two dehydrogenation reaction zones is essentially the quantity of feed to achieve the desired capacity increase. For example, when 33 percent of the total hydrocarbon feed is bypassed, there is a resulting 50 percent increase in total plant capacity when starting with two reaction zones and going to three reaction zones. In performing the overall process in this manner, there is no change in the steam flow rate and temperature to the steam superheater and the combined feed to the existing first dehydrogenation reaction zone remains unchanged from the original design, so that the critical pieces of equipment i.e., the steam superheater and two existing reactors are not directly affected by the capacity expansion. Capacity increases of 50-60 percent are feasible using the process of the present invention.

BRIEF DESCRIPTION OF THE DRAWING

[0007] The drawing is a simplified process flow diagram of a preferred embodiment of the present invention. The drawing is intended to be schematically illustrative of the present invention and is not a limitation thereof.

DETAILED DESCRIPTION OF THE INVENTION

[0008] Processes for the dehydrogenation of aromatic hydrocarbons are in a widespread commercial use. For instance, large quantities of styrene are produced by the dehydrogenation of efhylbenzene. The resultant styrene may be polymerized with itself or it may be copolymerized with butadiene, isoprene, acrylonitrile, etc. Other hydrocarbons which may be dehydrogenated in much the same manner include diethylbenzene, ethyl toluene, 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. The process can therefore be applied to the dehydrogenation of propane, butanes, hexanes or nonanes. However, since the great majority of the present commercial dehydrogenation processes are employed for the dehydrogenation of ethylbenzene, the following description of the subject invention will be presented primarily in terms of the dehydrogenation of ethylbenzene. This is not intended to exclude from the scope of the subject invention those alkylaromatic and acyclic hydrocarbons set out above or those having different ring structures including bicyclic compounds.

[0009] The dehydrogenation reaction is highly endothermic. Therefore, passing the reactants through a dehydrogenation catalyst bed results in a decrease in the reactant temperature. The endothermicity of the reaction is such that the temperature decrease removes the reactants from the desired temperature range. The reactants are actually cooled to such an extent that the desired reaction does not progress any further at a commercially feasible rate. The desired or commercially necessary per pass conversion therefore cannot be achieved by simply passing the reactants into contact with a single bed of dehydrogenation catalyst. For this reason, it has become standard commercial practice to in some manner perform interstage reheating. In interstage reheating the reactant effluent of a first bed of catalyst is heated to the desired inlet temperature of a second downstream bed of catalyst. This reheating can be performed through direct heat exchange as by the admixture of high temperature steam into the reactant stream emerging from the first catalyst bed.

[0010] A preferred method of interstage reheating comprises the use of indirect heat exchange. In this method, the effluent from a dehydrogenation zone is passed through a heat exchanger in which it is heated, and the reactants are then passed into the subsequent dehydrogenation zone. The high temperature fluid employed in this indirect heat exchange method may be high temperature steam, combustion gases, a high temperature process stream or other readily available high temperature fluids.

[0011] In accordance with the present invention, a first portion of the dehydrogenatable hydrocarbon feedstock is heated and introduced, preferably with steam, into a first dehydrogenation reaction zone operated without oxidative reheat to produce an effluent stream which is re-heated and introduced into a second dehydrogenation reaction zone. The resulting effluent from the second dehydrogenation reaction zone contains hydrogen produced during dehydrogenation which is then available for catalytic oxidation to produce heat to reheat the reactants before being passed into the downstream dehydrogenation catalyst. The resulting effluent from the second dehydrogenation reaction zone, a second portion of the dehydrogenatable hydrocarbon feedstock, oxygen, and optionally, steam are reacted in a third dehydrogenation reaction zone operated with oxidative reheat.

[0012] The second portion of the dehydrogenatable hydrocarbon feedstock, oxygen and steam, if present, are preferably admixed before entering the third dehydrogenation reaction zone in order to ensure a homogeneous feed to the oxidative reheat catalyst in order to achieve the desired reaction temperature before contacting the dehydrogenation catalyst. This homogeneity also eliminates the possibility of having explosive concentrations of hydrocarbons.

[0013] The incentive for employing oxidative reheat is the recognition that the combustion of the hydrogen generated in the dehydrogenation reaction zones performs two functions which are beneficial in the dehydrogenation process. First, the consumption of the hydrogen is beneficial in shifting the equilibrium of the dehydrogenation reaction to favor increased amounts of dehydrogenation. Second, the selective combustion of the hydrogen will release heat sufficient to reheat the reactants to the desired dehydrogenation conditions.

[0014] The oxidation is preferably accomplished in the presence of a catalyst which selectively promotes the oxidation of hydrogen as compared to the destructive combustion or oxidation of the more valuable feed and product hydrocarbons. The selective combustion method of interstage reheating presents a more economical dehydrogenation process.

[0015] Despite the advances which have been achieved in the arts of catalysis and hydrocarbon conversion, the ultimate conversion which can be achieved during a single passage through a dehydrogenation zone is limited to an amount less than total conversion. That is, it is impossible to achieve a 100% conversion of a feed hydrocarbon to a corresponding product dehydrogenated hydrocarbon. A basic limitation in the degree of conversion which may be achieved in any dehydrogenation processes is the equilibrium concentration of the various reactants at the temperatures employed. The effluent stream of a catalytic dehydrogenation zone will therefore comprise a mixture of the feed hydrocarbon, the dehydrogenated hydrocarbon product, and hydrogen. Generally, it is necessary to separate and recover the dehydrogenated hydrocarbon product and to recycle the unconverted feed hydrocarbon. The greater the rate of conversion which is achieved in the dehydrogenation zone, the smaller the amount of unconverted material is realized which must be recycled. The separation of the product and unreacted hydrocarbons requires extensive capital equipment and consumes large amounts of utilities in the form of heat and electrical power. It is therefore desirable to increase the conversion which is achieved per pass in the dehydrogenation zone and to thereby decrease the amount of material which must be separated and recycled. A higher per passage conversion will also allow a smaller reaction zone to be employed in the process with the associated reduction in the cost of the reactors, catalyst and utilities cost of operating the reaction zone. For these reasons, it is highly desirable to achieve increased rates of total conversion during the passage of the dehydrogenation zone feed stream through a multibed dehydrogenation zone.

[0016] In the oxidative reheat process, an oxygen-containing gas stream is preferably admixed with the effluent of a preceding dehydrogenation reaction zone and the resulting admixture along with a portion of the dehydrogenatable hydrocarbon feedstock is passed into a bed of selective hydrogen oxidation catalyst. To achieve the optimum levels of performance and safety in this process, it is necessary to closely control the rate at which oxygen is passed into the process in this manner.

[0017] An insufficient amount of oxygen will result in a less than desired consumption of hydrogen and more importantly a less than desired reheating of the reactant stream. The result will be a decrease in the degree of dehydrogenation achieved during passage through the overall reaction zone. It is not normally desired to inject an excess amount of oxygen into any part of the dehydrogenation zone above that required to perform the desired degree of hydrogen combustion.

[0018] The passage of an excess amount of oxygen into the dehydrogenation zone will also have detrimental effects upon the long term operation of the process. For instance, oxygen will normally serve to deactivate or poison some commercially employed dehydrogenation catalyst. It is therefore undesirable to have residual oxygen emerging from the oxidation catalyst bed and thereupon contacting dehydrogenation catalyst. Operation of the dehydrogenation zone in a manner which does not result in the total consumption of the oxygen is also undesirable because of the obvious explosive nature of oxygen-hydrocarbon mixtures. The explosive nature of these mixtures can, however, be essentially negated by properly operating the process to avoid the presence of mixtures being within the explosive range, as through the use of diluents and intentionally low oxygen addition rates, and the presence of a sufficient amount of solid material to act as an explosion suppression means. Lastly, the presence of oxygen is not normally desired in vessels containing hydrocarbons as the oxygen may react with the hydrocarbons to form various undesired oxygenated compounds. [0019] The structure of the dehydrogenation reaction zones may be varied by changing the type of catalyst bed which is employed. For instance, radial flow through annular catalyst beds as well as vertical flow through cylindrical catalyst beds. It is to be noted that with radial flow, the beds of dehydrogenation catalyst and oxidation catalyst may be concentrically located at the same elevation within the vessel or vessels. Either the oxidation catalyst or the dehydrogenation catalyst may be located in the outer bed of this arrangement. The gas flow would then pass through cylindrical center pipe regions located in the middle of the radial flow catalyst beds and through annular gas collection and distribution void volumes located between the outer surface of the catalyst beds and the inner wall of the vessel. Variation is also possible in the number of beds of catalyst which may be employed. Suitable systems for catalyst deployment may be patterned after those presented in US 3,498,755; US 3,515,763; and US 3,751,232.

[0020] Dehydrogenation catalysts generally consist of one or more metallic components selected from Groups VI and VIII of the Periodic Table. One typical catalyst for the dehydrogenation of alkylaromatics comprises 85 wt% ferric oxide, 2 wt% chromia, 12 wt% potassium hydroxide and 1 wt% sodium hydroxide. A second dehydrogenation catalyst, which is used commercially, consists of 87 to 90 wt% ferric oxide, 2 to 3 wt% chromium oxide and from 8 to 10 wt% potassium oxide. A third typical 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 demonstrated by the teachings of US 3,387,053, which describes the manufacture of a catalytic composite of at least 35 wt% iron oxide as an active catalytic agent, from 1 to 8 wt% zinc or copper oxide, 0.5 to 50 wt% of an alkali promoter, and from 1 to 5 wt% chromic oxide as a stabilizer and a binding agent. US 4,467,046 also describes a catalyst for the dehydrogenation of ethylbenzene in the presence of steam. This catalyst contains 15 to 30 wt% potassium oxide, 2 to 8 wt% cerium oxide, 1.5 to 6 wt% molybdenum oxide, 1 to 4 wt% calcium carbonate, with the balance being iron oxide.

[0021] Dehydrogenation conditions in general include a temperature of 500° to 750⁰C and preferably 565° to 675°C. The temperature required for efficient operation of any specific dehydrogenation process will depend on the feed hydrocarbon and the activity of the catalyst employed. The pressure maintained within the dehydrogenation zone may range from 100 to 750 mm Hg, with a preferred range of pressures being from 250 to 700 mm Hg. The operating pressure within the dehydrogenation zone is measured at the inlet, midsection, and outlet of the zone to thereby provide an average pressure. The feed stream is charged to the dehydrogenation zone at a liquid hourly space velocity from 0.1 to 2.0 hr ^"', and preferably from 0.1 to 1.0 hr^"1, based on the total liquid hydrocarbon charged at 15.6°C.

[0022] The hydrocarbon feed to be dehydrogenated is preferably admixed with superheated steam to counteract the temperature lowering effect of the endothermic dehydrogenation reaction. The presence of steam has also been described as benefiting the stability of the dehydrogenation catalyst by preventing the accumulation of carbon deposits. Preferably, the steam is admixed with the other components of the feed stream at a rate of 0.5 to 1.5 pound of steam per pound of feed hydrocarbon. Other quantities of steam may be added after one or more subsequent dehydrogenation catalyst beds if desired. However, the dehydrogenation zone effluent stream should contain less than 3 pounds of steam per pound of product hydrocarbon and preferably less than 2 pounds of steam per pound of product hydrocarbon.

[0023] The vaporous effluent stream from the last dehydrogenation zone may be heat exchanged against a stream of steam, a reactant stream of this or another process, or used as a heat source for fractionation. Commercially, 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 subject to the constraints set out above. Preferably, the cooling is sufficient to condense at least 95 mole percent of the C6+ hydrocarbons, i.e. hydrocarbons having 6 or more carbon atoms per molecule, and at least 95 mole percent of the water vapor in the dehydrogenation zone effluent stream. Essentially all of the dehydrogenated hydrocarbon product such as styrene, most water and other readily condensable compounds present in the effluent stream are thereby converted to liquids. This produces a mixed phase stream which is passed into a phase separation vessel. This procedure allows the facile crude separation by decantation of the hydrocarbons from the water and hydrogen present in the effluent stream. The dehydrogenated hydrocarbon product present in the dehydrogenation zone effluent stream becomes part of the hydrocarbon stream which is withdrawn from the separation vessel and transferred to the proper separation facilities. The dehydrogenated hydrocarbon product is preferably recovered from the hydrocarbon stream by using one of the several fractionation systems known in the art. This fractionation will preferably yield a relatively pure stream of the unconverted hydrocarbon feed such as ethylbenzene, which may be recycled for improved economics. An additional hydrocarbon stream comprising by-products of the dehydrogenation reaction may also be obtained during product fractionation. For example, in the production of styrene from ethylbenzene, benzene and toluene may be recovered, and may be recycled in part as taught in US 3,409,689 and GB 1,238,602 or entirely rejected from the process. If desired, methods other than fractionation may be used to recover the dehydrogenated hydrocarbon product. For instance, US 3,784,620 teaches the separation of styrene and ethylbenzene through the use of a polyamide permeation membrane such as nylon-6 and nylon 6, 10. US 3,513,213 teaches a separation method employing liquid-liquid extraction in which anhydrous silver fluoroborate is used as the solvent. Similar separation methods using cuprous fluoroborates and cuprous fluorophosphates are described in US 3,517,079; US 3,517,080; and US 3,517,081. [0024] The oxygen supply stream to the process may be air but is preferably a gas having a higher oxygen content than air. It is preferred that the oxygen supply stream has a nitrogen content less than 10 mole percent, with the use of substantially pure oxygen being highly preferred if it is economically viable. The preferred oxygen concentration in the oxygen supply stream is primarily a matter of economics and may be determined by a comparison of the advantage of having pure oxygen to the cost of obtaining the oxygen. The basic disadvantages of the presence of nitrogen are the dilution of the hydrogen-containing gas stream removed from the product separation vessel and the fact that the nitrogen passes through the dehydrogenation zone thereby increasing the pressure drop through the catalyst bed and the absolute pressure being maintained within the dehydrogenation zone. On the other hand, the presence of nitrogen favorably affects the equilibrium conversion level by acting as a diluent.

[0025] The oxidation catalyst employed in the oxidative reheat or oxidation zone to promote the hydrogen oxidation may be any commercially suitable catalyst. The oxidation catalyst will have a different composition than the dehydrogenation catalyst. Preferably, the oxidation catalyst will have a high selectivity for the oxidation of hydrogen with only small amounts of the feed or product hydrocarbons being oxidized. A preferred oxidation catalyst comprises an IUPAC Group 7, 8, or 9 noble metal and at least one other metal or metal cation with both of these materials being present in small amounts on a refractory solid support. The preferred noble metals are platinum and palladium, but the use of ruthenium, rhodium, osmium and indium is also contemplated. In an embodiment, the noble metal is present in an amount ranging from 0.01 to 5.0 wt% of the finished catalyst. The metal or metal cation is preferably chosen from IUPAC Groups I or 2 and is present in an amount ranging from 0.01 to 20 wt% of the finished 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 IUPAC Group 14.

[0026] In a preferred embodiment, the refractory solid support of the oxidation catalyst is alumina having a surface area between 1 and 300 m²/g; an apparent bulk density between 0.2 and 1.5 g/cc; and an average pore size greater than 20 angstroms. The metal-containing components are preferably impregnated into solid particles of the solid support by immersion in an aqueous solution followed by drying and calcination at a temperature ranging from 500⁰C to 1200⁰C in air. The support may be in the form of spheres, pellets or extrudates. The total amount of oxidation catalyst present within the dehydrogenation zone is preferably less than 30 wt% of the total amount of dehydrogenation catalyst and more preferably is between 5 and 15 wt% of this total amount of dehydrogenation catalyst.

[0027] The conditions utilized during the contacting of the reactant stream with the bed of oxidation catalyst will be set to a large extent by the previously described dehydrogenation conditions. The preferred outlet temperature of the oxidation catalyst is the preferred inlet of the downstream bed of dehydrogenation catalyst. The temperature rise across the oxidation catalyst is preferably adjusted by the amount of hydrogen conversion across the oxidation catalyst. The liquid hourly space velocity, based on the liquid hydrocarbon charge at 15.6°C, is preferably between 2 and 20 hr^"1.

DETAILED DESCRIPTION OF THE DRAWING

[0028] In the drawing, the process of the present invention is illustrated by means of a simplified schematic flow diagram in which such details as pumps, instrumentation, heat- exchange and heat-recovery circuits, compressors and similar hardware have been deleted as being non-essential to an understanding of the techniques involved. The use of such miscellaneous equipment is well within the purview of one skilled in the art.

[0029] Referring now to the drawing, a hydrocarbon feed stream comprising a C3+ feed hydrocarbon, i.e. a hydrocarbon having 3 or more carbon atoms per molecule, is introduced into the process via line 1 and bifurcated to provide a first portion and a second portion. The first portion of the feed stream is carried via line 2 and is combined with steam provided by line 3 and the resulting mixture is carried via line 4 and introduced into dehydrogenation reaction zone 5. Dehydrogenation zone 5 is operated without oxidative reheating and the resulting effluent is carried via line 6, heated via indirect heat exchange (not shown) and introduced into dehydrogenation reaction zone 7. Dehydrogenation zone 7 is operated without oxidative reheating and the resulting effluent is transported via line 8 and is combined with the second portion of the feed stream which is carried via lines 13 and 9. The resulting effluent from dehydrogenation zone 7 carried via line 8 is also combined with an oxygen and steam mixture provided via lines 14 and 9. This resulting mixture is carried via line 10 and introduced into dehydrogenation reaction zone 11 which is conducted with oxidative reheating. Dehydrogenation reaction zone 11 contains an oxidation zone 16 and a dehydrogenation zone 17. The resulting effluent from dehydrogenation reaction zone 11 comprises the product dehydrogenated hydrocarbon which is carried via line 12 and recovered.

Claims

CLAIMS:

1. A process for the catalytic dehydrogenation of a C3+ feed hydrocarbon comprising:

(a) passing a first portion of a feed stream comprising the C3+ feed hydrocarbon through a first bed of dehydrogenation catalyst at dehydrogenation conditions in a first dehydrogenation zone and producing a first dehydrogenation zone effluent stream comprising hydrogen, the C3+ feed hydrocarbon and a C3+ product hydrocarbon;

(b) heating and passing at least a portion of the first dehydrogenation zone effluent stream through a second bed of dehydrogenation catalyst at dehydrogenation conditions in a second dehydrogenation zone and producing a second dehydrogenation zone effluent stream comprising hydrogen, C3+ feed hydrocarbon and C3+ product hydrocarbon;

(c) passing at least a portion of the second dehydrogenation zone effluent stream, a second portion of the feed stream comprising the C3+ feed hydrocarbon and oxygen into a separate bed of selective hydrogen oxidation catalyst at oxidation conditions in an oxidation zone to produce an oxidation zone effluent; (d) passing at least a portion of the oxidation zone effluent through a third bed of dehydrogenation catalyst at dehydrogenation conditions in a third dehydrogenation zone to produce a third dehydrogenation zone effluent stream which comprises the product hydrocarbon; and

(e) recovering the product hydrocarbon.

2. The process of claim 1 further comprising mixing steam with the first portion of the feed stream before passing the first portion of the feed stream and steam through the first bed of dehydrogenation catalyst.

3. The process of claim 2 wherein the amount of steam ranges from 0.5 to 1.5 pounds per pound of C3+ feed hydrocarbon in the first portion of the feed stream.

4. The process of any one of claims 1 or 2 wherein steam is mixed with the portion of the second dehydrogenation zone effluent stream, the second portion of the feed stream, and oxygen before passing the portion of the second dehydrogenation zone effluent stream, the second portion of the feed stream, oxygen, and steam into the separate bed of selective hydrogen oxidation catalyst.

5. The process of any one of claims 1 or 2 wherein the C3+ feed hydrocarbon is an alkylaromatic hydrocarbon.

6. The process of any one of claims 1 or 2 wherein the C3+ feed hydrocarbon is ethylbenzene.

7. The process of any one of claims 1 or 2 wherein the C3+ feed hydrocarbon is propane.

8. The process of any one of claims 1 or 2 wherein the C3+ feed hydrocarbon is butane.

9. The process of any one of claims 1 or 2 wherein the dehydrogenation conditions include a temperature from 500⁰C to 750⁰C and a pressure from 100 to 750 mm Hg.

10. The process of claim 9 wherein the dehydrogenation conditions include a pressure from 250 to 700 mm Hg.