(en)A process for the production and recovery of mono-olefinic hydrocarbons. A paraffinic hydrocarbon feed stream is contacted with a dehydrogenation catalyst to produce olefinic hydrocarbons with minor side-reaction product comprising primarily aromatic type hydrocarbons. The effluent material from the dehydrogenation catalyst is passed into a separation zone wherein the olefinic and aromatic hydrocarbons are separated from the unreacted paraffinic hydrocarbons. The unreacted paraffinic hydrocarbons are recycled to the dehydrogenation reactor for further production of olefinic hydrocarbons while the olefinic hydrocarbons are recovered as a relatively pure hydrocarbon extract stream.

1.ApplicationNumber: US-3617504D-A
1.PublishNumber: US-3617504-A
2.Date Publish: 19711102
3.Inventor: ROY C. BERG

4.Inventor Harmonized: BERG ROY C()

5.Country: US
6.Claims:
(en)A process for the production and recovery of mono-olefinic hydrocarbons. A paraffinic hydrocarbon feed stream is contacted with a dehydrogenation catalyst to produce olefinic hydrocarbons with minor side-reaction product comprising primarily aromatic type hydrocarbons. The effluent material from the dehydrogenation catalyst is passed into a separation zone wherein the olefinic and aromatic hydrocarbons are separated from the unreacted paraffinic hydrocarbons. The unreacted paraffinic hydrocarbons are recycled to the dehydrogenation reactor for further production of olefinic hydrocarbons while the olefinic hydrocarbons are recovered as a relatively pure hydrocarbon extract stream.

7.Description:
(en)United States Patent [72] inventor Roy C. Berg Park Ridge, Ill. 21 1 Appl. No. 859,737 [22] Filed Sept. 22, 1969 [45] Patented Nov. 2, 1971 [73] Assignee Universal Oil Products Company Des Plaines, Ill.
[54] PRODUCTION AND RECOVERY OF OLEFINIC HYDROCARBONS 27 Claims, 2 Drawing Figs.
[52] U.S. Cl 208/100, 208/102, 208/310, 260/683.3 [5 1] Int. Cl. C07c 5/18 [50] 208/310, 102, 100, 99; 260/676 MS, 683.3
[5 6] References Cited UNITED STATES PATENTS 2,950,336 8/1960 Kimberlin et a]. 208/310 3,063,934 11/1962 Epperbyetal 260/676 3,437,585 4/1969 Kuchar 260/6833 3,510,423 5/1970 Neuziletal 260/676 Primary Examinerl-1erbert Levine Attorneys-James R. Hoatson, Jr. and Robert W. Erickson Over/mad Fraslr Feed Freer/00a flan Adsorpfin -$apararian 1 Extract 1 PRODUCTION AND RECOVERY OF OLEFINIC HYDROCARBONS BACKGROUND OF THE INVENTION 1. Field of the Invention This invention relates to a combination process for the production and recovery of mono-oletinic hydrocarbons. More specifically, the combination invention relates to the dehydrogenation of paraffinic hydrocarbons to form primarily mono-oletins followed by the separation of the olefinic hydrocarbons from the remaining parafiinic feed through the use of a crystalline aluminosilicate adsorbent, which selectively excludes the paraffinic hydrocarbons, The parafiinic hydrocarbons are recycled to the dehydrogenation reaction zone for further reaction to mono-olefins.
2. Description of the Prior Art Recently, attention within the chemical and petroleum industry has been focused upon the problem of acquiring long chain mono-olefinic hydrocarbons. in particular, a substantial demand has been established for normal mono-olefins having from about 6 to about 20 carbon atoms per molecule which are valuable raw materials for use in the production of alkenyl sulfonates or alkylaromatic hydrocarbons which can be converted into biodegradable detergents. The production of biodegradable detergents has been emphasized in this and most other industrialized countries to reduce the effects of water pollution which, in part, is caused by the branched chain alkylaromatic-type detergents which are not susceptible to bacterial degradation when dispersed into natural waters with raw or semitreated sewage.
in response to this demand for relatively straight-chain mono-olefins, the art has developed dehydrogenation catalysts which have high degrees of selectivity for the production of normal mono-olefinic hydrocarbons from feed streams comprising normal paraffinic hydrocarbons. One particular method for the production of relatively long straight-chain monoolefinic hydrocarbons has been the selective dehydrogenation of a normal paraffin feed stream by contacting it and hydrogen with a nonacidic, alumina supported, platinum metal-containing catalyst. The principal feature of using this particular type catalyst in a process for the direct dehydrogenation of long-chain normal paraffins, involves the capability of this catalyst to sustain a high level of selectivity for the production of the desired normal mono-olefinic hydrocarbons with the additional ability to suppress undesirable side or secondary reactions such as isomerization, secondary dehydrogenation, dehydrocycloization, polymerization and cracking.
In direct dehydrogenation methods employing the general class of catalysts as described above, economic considerations require that the unreacted parafiinic hydrocarbons be recycled to the dehydrogenation zone after separation of the olefinic hydrocarbons from the paraffinic hydrocarbons and that the paraftinic recycle stream be reacted to extinctionthat is, substantially all the paraftinic hydrocarbons fed to the dehydrogenation unit be eventually converted into primarily mono-oletins. in order to reduce the aforementioned undesirable side reactions, it is generally recognized that relatively low severity dehydrogenation conditions should be maintained. The obvious consequence of the relatively low severity operating conditions occurring in the direct dehydrogenation process is a low conversion of the paraffinic material passed into the dehydrogenation zone to monoolefinic hydrocarbons. The low conversions of paraffinic feed yields a dehydrogenation process effluent containing relatively low concentrations of the product mono-olefinic hydrocarbon which generally are nonseparable from the parafiinic hydrocarbon by present day fractionation techniques. In cases where the relatively long chain mono-olefin is to be used in the production of biodegradable detergent, the unreacted paraffinic material in the dehydrogenation reaction zone effluent can be ,passed into an alkylation zone in admixture with the mono-olefinic hydrocarbons wherein the more reactive olefinic material is alkylated while the less reactive paraflinic material passes through unreacted. The subsequent mixture of alkylaromatic hydrocarbons and the unreacted paraffinic feed present in the alkylation zone effluent is then separated by fractionation techniques and the paraftinic hydrocarbon can be recycled to the dehydrogenation zone for further reaction to mono-olefinic hydrocarbons.
I have found that when employing the general class of catalysts as recited above in the direct dehydrogenation reaction zone that at conversion levels in the range of from about 5 to about 30 percent that there is an inherency present within the process that does not allow the complete elimination of undesirable side reaction products. It is believed that the presence of the platinum group metal because of its wellrecognized catalytic ability as to isomerization, dehydrogenation, dehydrocyclization, etc., is primarily responsible for the inability to totally eliminate the occurrence of concurrent side reaction products from being produced in the dehydrogenation zone. The relatively long chain mono-olefinic hydrocarbons produced in the direct dehydrogenation zone boil roughly at the same temperatures as the unreacted parafiinic hydrocarbons and consequently are not easily separable from the parafiins. Consequently, in most cases where the long chain mono-olefins are produced, the effluent stream from the direct dehydrogenation zone is passed into a reaction zone wherein the mono-olefins contained therein can be reacted to form products such as the aforementioned alkylsromatic hydrocarbons, alcohols, ketones, esters, etc. The unreacted paraftins are recycled back to the dehydrogenation zone for further contact with the catalysts contained therein and subsequent production of additional amounts of mono-olefinic hydrocarbons. Accordingly, any side reaction products that are synthesized in the dehydrogenation process plus any contaminants present in the feed stream which boil within the boiling range of the recycle paraffin stream can accumulate in this recycle paraffin stream. This contamination of the recycled paraffin stream leads to several significant adverse effects: (1) these contaminants are somewhat refractory to the conditions utilized in the dehydrogenation step, and consequently in order to maintain a constant weight percent conversion of this stream to normal olefms, it is necessary to raise the severity level utilized in the dehydrogenation step; and (2) the presence of these contaminants greatly increases the rate of formation of deleterious carbonaceous deposits on the catalyst. Both of these effects eventually result in degradation of the high selectivity of the dehydrogenation catalyst. l have found that the principal contaminants causing this catalyst degradation problem are aromatic-type contaminants and that this preferred dehydrogenationprocess can be significantly improved by removing at least a portion or all of the aforementioned contaminants, and thus prevent the buildup of these contaminants in the recycle paraffin stream.
It has been recognized in the prior art processes involving direct dehydrogenation of normal parafiins wherein a majority of the output or effluent stream from the dehydrogenation zone is recycled thereto for further reaction to produce additional mono-olefinic materials that the contaminant buildup primarily comprising aromatic-type hydrocarbons can be prevented by two general methods which generally include: (1) acid treating of a portion or all of the recycle paraffin stream to remove the aromatic contaminants present, therein reducing contaminant buildup and additionally improving and maintaining stable dehydrogenation catalyst operations; and (2) the removal via a slip or drag stream of a portion of the recycle paraffin stream.
The acid washing techniques generally remove substantially all the aromatic hydrocarbons from the recycle paraftin stream and allow removal of these aromatic hydrocarbons from the parafiinic material which is recycled back to the dehydrogenation zone. The aromatic contaminants generally and up as the preferably soluble components present in the acid phase and can easily be recovered for removal from the process. The techniques employing slip or drag streams generally involve removing a quantity of the paraffinic recycle material from the process at a rate in which the aromatic material removed from the slip or drag stream is generally equal to the quantity of the aromau'c side reaction products produced in the dehydrogenation zone. Of course, this particular type operation reduces the overall efficiency of the dehydrogenation process in that the paraffinic slip or drag stream removed reduces the actual yield based on fresh feed passed into the dehydrogenation zone reactor. The recycle paraffin material removed as a slip stream can then be acid treated and recovered with the parafiinic material being returned to the dehydrogenation process to reduce yield losses in the process. Both of the aforementioned methods for removal of the aromatic contaminants which boil within the boiling range of parafiinic recycle stock to the dehydrogenation zone require that an olefinic-parafiinic effluent stream from the dehydrogenation zone be passed into a subsequent reaction zone wherein the mono-olefinic hydrocarbons can be converted, by well-known processes and in the presence of paraffins to a desired end product requiring as one of its raw material the mono-olefinic hydrocarbon. It can readily be seen that in this particular type of operations wherein the total effluent from the dehydrogenation zone which comprises primarily paraffins and a relatively small quantity of olefins, that large quantities of material must be handled in order to effectively utilize the low conversions which necessarily are employed in the dehydrogenation zone.
I have found that by incorporatingand adsorptive separation-type process which employs a crystalline aluminosilicate adsorbent that the buildup of aromatic contaminants in the paraffin recycle stream is eliminated while producing a concentrated and relatively pure mono-olefinic process stream. In using the combination process of this invention-the direct dehydrogenation reaction zone in combination with an adsorptive separation process employing a crystalline aluminosilicate adsorbent-the deleterious aromatic contaminants which are inevitably produced in the direct dehydrogenation zone using a dehydrogenation catalyst of a class mentioned previously are removed from the process, as are the mono-olef'mic hydrocarbons. While the combination of the process of this invention does not reduce the total quantity of parafiin recycle material handled, it does eliminate the need for chemical separation of the mono-olefins from the unreacted paraffinic hydrocarbons which are withdrawn from the dehydrogenation zone while additionally yielding a highly concentrated olefin stream and reducing aromatic buildup in the recycle paraffin stream.
SUMMARY OF THE INVENTION It is an object of the present process to provide a combination process for: l the selective catalytic dehydrogenation of a hydrocarbon stream containing normal paraffinic hydrocarbons wherein unreacted normal parafiin hydrocarbons are recovered and recycled to extinction; and (2) the separation of the effluent material comprising paraftinic and oletlnic hydrocarbons in an adsorptive-separation process wherein the unreacted normal paraffinic hydrocarbons are recycled to the dehydrogenation zone and a stream comprising the monoolefinic hydrocarbons are adsorbed and subsequently recovered from said adsorptive separation process. A second and more specific object relates to a selective catalytic dehydrogenation process in combination with an adsorptive separation process wherein said dehydrogenation process employs a nonacidic, alumina supported, platinum metal containing catalyst to dehydrogenate a hydrocarbon stream containing normal paraffin hydrocarbons and an adsorptive-separation process employing a crystalline aluminosilicate having uniform pore openings of about 6 through about angstrom units and containing exchangeable cations selected from the group consisting of Group IA, Group "B and Group "A metals wherein said effluent stream from said dehydrogenation process is passed into said adsorptive-separation process at conditions to effect the selective adsorption of olefinic hydrocarbons and substantially all side reaction aromatic hydrocarbons within said adsorbent while selectively rejecting the paraffinic hydrocarbons which are recycled to the dehydrogenation zone to effect their reaction to extinction. The selectively adsorbed material is removed from said adsorbent as a concentrated stream comprising olefinic hydrocarbons and a relatively small portion of aromatic hydrocarbons.
Hydrocarbon feed streams which can be charged to the present combination invention contain normal paraffin hydrocarbons having at least six carbon atoms per molecule and especially normal paraffin hydrocarbons having from about nine to about 20 carbon atoms per molecule. Representative members of this class are; hexane, heptane, nonane, decane, undecane, dodecane, tridecane, tetradecane, pentadecane, etc. and mixtures thereof. Hydrocarbon feed streams containing normal paraffins of about ten to 18 carbon atoms per molecule are particularly advantageously used since they produce normal mono-olefins which can be used to produce sulfonated arylalkane-type detergents having superior biodegradability. A feed mixture containing a four or five homolog spread provides an excellent charge stock. it is preferred that the amount of non normal hydrocarbons present in this hydrocarbon feed stream be kept at low levels. Preferably, the nonnormal hydrocarbons present in hydrocar bon feed streams to the dehydrogenation zone should be in concentrations of less than about l0 wt. percent of the normal paraffin hydrocarbons fed to the dehydrogenation zone. In accordance with the present invention, it is preferred that the amount of aromatic contaminants present in the dehydrogenation zone feed stream be held to substantially negligible levels--that is, less than about 1 wt. percent and preferably less than about 0.5 wt. percent of the feed stream. Typically, the aromatic contaminants present in the feed stream generally are from the class of compounds consisting of alkylaromatics, alkylindanes, and bicyclic aromatics. Accordingly,.it is within the scope of the present invention to treat the dehydrogenation zone hydrocarbon feed stream to remove substantially all of the aromatic contaminants present prior to passage into the dehydrogenation zone.
The catalyst used in the dehydrogenation step of the combination invention is specifically designed to effect dehydrogenation of normal paraffin hydrocarbons while suppressing side reactions. The preferred catalysts comprise a platinum group component and an alkali component composited with an alumina carrier material. Although it is not essential, it is generally preferred that this catalyst also contain a component selected from the group consisting of arsenic, bismuth, antimony, sulfur, selenium, tellurium, rhenium, germanium, or mixtures thereof.
The aluminum carrier material of this dehydrogenation catalyst generally has an apparent bulk density of less than about 0.50 gm./cc. with the lower limit from about 0.15
gm/cc. The surface area characteristics of this alumina are preferably such that the average pore diameter is about 20 to about 300 angstrom units, pore volume is about 0.10 to about 1.0 mm./gm., and the surface area is about l00 to about 700 M lgm. Particularly preferred alumina carriers comprise gamma-alumina. The alumina carrier material may be manufactured by any suitable method including well-known manufacturing procedures as detailed in U.S. Pat. No. 2,620,314.
The alkali component is selected from both the alkali metals and the alkali earth metals with best results obtained with lithium. Preferably, the catalyst containing this component in the amount of from 0.01 to about 3.0 wt. percent. Additionally, this component is preferably combined with alumina by impregnation.
The platinum group component is generally selected from the group of palladium, iridium, lithium, rhodium, osmium and platinum, with platinum giving the best results. The platinum group component may be used in the form of the elemental metal or as a suitable compound such as dioxide, sulfide, halide, etc., although it is generally preferred that it be used in a concentration calculated on an elemental basis of from about 0.05 wt. percent to about 5.0 wt. percent of the catalytic composite with best results obtained at a level of from about 0.2 to about 1.5 wt. percent. This component may be composited in any suitable manner with impregnation by water soluble compounds such as chloroplatinic acid being especially preferred.
Preferably, the dehydrogenation catalyst contains a fourth component selected from the group consisting of arsenic, antimony, bismuth, sulfur, selenium, tellurium, rhenium, germanium or mixtures thereof. Arsenic is particularly preferred. This component is preferably used with good results in an amount of from about 0.01 percent to about 1.0 percent by weight of the final composite. Moreover, this component is preferably present in an atomic ratio to the platinum component of from about 0.1 to about 0.8 with intermediate concentrations of about 0.2 to about 0.5 being highly effective. This component can be composited in any suitable manner-a particularly preferred method employing an impregnation solution of water soluble compounds such as arsenic pentoxide.
The adsorbent used in the adsorptive separation zone is specifically designed to effectuate the separation of normal paraffin hydrocarbons from mono-olefinic hydrocarbons and aromatic hydrocarbons produced by undesirable side reactions occurring in the dehydrogenation zone.
. Specifically, the adsorbent comprises a crystalline aluminosilicate zeolite containing selected cations at the exchangeable cationic sites to effectuate the aforementioned separation. The crystalline aluminosilicate zeolites can be described as a three dimensional framework of SiO and A tetrahedra in which the tetrahedra are cross-linked by the sharing of oxygen atoms whereby the ratio of the total alumina and silicon atoms to oxygen atoms is 1:2. In their hydrated form, the tetrahedra are occupied by water molecules. Dehydration results in crystals interlaced with the channels of molecular dimensions. In their hydrated form, the aluminosilicates may be represented by the formula represented in equation 1 below,
where M is a cation which balances the electrovalence of the tetrahedra, n represents the valence of the cation, w the moles of Sill, and y the moles of water. M, as is well known in the art, is generally referred to as representing the exchangeable cations present in the crystalline aluminosilicate and may be any number of metal ions depending on whether the aluminosilicate is synthesized or occurs naturally. Typical cations which can be present at the cation exchangeable sites within the crystalline aluminosilicate include selected cations from the Group IA, Group IIA, Group IB, Group 118 metals and mixtures thereof. The preferred cations include sodium, silver, copper and potassium or combinations thereof. These particular cations display superior selectivity for olefin adsorption and in most cases do not cause excessive olefin polymerization or isomerization to adversely effect the adsorption process. Although the proportions of inorganic oxides in the silicates in their spacial arrangement may slightly vary, effecting distinct properties in the aluminosilicates, the two main characteristics of these materials is the presence of at least 0.5 equivalent of an ion of positive valence per gram atom of alumina, and an ability to undergo dehydration without substantially effecting the tetrahedra framework. In this respect, these characteristics are essential for obtaining selective separations in accordance with the combination process of this invention.
Representative materials of the above-mentioned general class of crystalline aluminosilicates include the synthesized crystalline aluminosilicates characterized by a three-dimensional array of pores which consist of cages interconnected by windows of from about 8 to about 9 angstrom units. Materials falling within the above-mentioned dimensional characterization include a general class of aluminosilicates comprising the faujasites. Included in the class of faujasites are the crystalline aluminosilicates characterized as a zeolite type X and a zeolite type Y and it is preferred to employ these type zeolites as the adsorbent material used within the adsorptive-separation unit in order to effect the required paraffin-olefin plus aromatic separation. The type X and type Y zeolites have been found to be the most capable crystalline aluminosilicates for the aforementioned separation and consequently, are the preferred adsorbents to be used in the combination process of this invention.
The zeolite type X can be represented in terms of lower ratios of oxides of silica and alumina as shown below in equation 2,
wherein M is a cation having a valence of not more than 3, n represents the valence of M, and y is a value of up to 8, depending on the identity of M and the degree of hydration of the crystal.
The zeolite type Y expressed in oxide mole ratios is shown in equation 3 below,
1.0i0.2M ,,,:Al 0 :wSi0,:yI-I 0 (3) wherein M is a cation and generally comprising sodium and n represents the valence of M, w is a value ranging from 3 to about 6 and y may be any value up to about 9.
In referring to a type X or type Y structured zeolite, the crystalline structure of three-dimensionally connected tetrahedra having 12 to 13 angstrom cages connected by 8 to 9 angstrom windows are presumed to be essentially unaltered and it is assumed when referring to the different type X and type Y structured zeolites that the essential difference between the individually mentioned structured zeolites is basically either the water content contained therein or the particular type or types of cations present at the exchangeable cationic sites within these zeolites, or both differences in water content and the types of exchangeable cations present at the exchangeable cationic sites within the zeolite.
The aromatic contaminants which contribute to the deactivation of the dehydrogenation catalysts generally comprise alkylaromatics, alkylindanes, and bicyclic aromatics. The concentration of the aromatics in the dehydrogenation zone effluent is generally in the 0.5 wt. percent range with approximately 20 percent of the aromatics comprising polynuclear aromatics of the class including naphthalenes, biphenols, or alkylnaphthalenes with the remaining aromatics comprising alkyl-substituted benzenes including di-isopropyl benzene, toluene, xylenes and other alkyl-substituted benzenes. The side reaction products produced in the dehydrogenation zone have been found to be adsorbed by certain crystalline aluminosilicate type adsorbents. Accordingly, it is within the scope of the combination process of this invention to employ a desorbent in the adsorptive-separation zone that, in addition to desorbing adsorbed olefinic hydrocarbons from the molecular sieves, also displaces adsorbed aromatic hydrocarbons thereby effecting a continuous removal of side reaction produced aromatic hydrocarbons which substantially reduces, and in most cases, totally eliminates any buildup of aromatic hydrocarbons in the paraffin recycle stream going to the dehydrogenation zone. This substantial reduction of aromatic hydrocarbons being recycled to the dehydrogenation zone via the paraffinic recycle stream enhances catalyst selectivity, conversion levels and activity and eliminates carbon buildup on the dehydrogenation zone catalyst.
The term selectivity when referring to the adsorptiveseparation zone refers to the ratio of the concentrations of two components present within the adsorbent and externally surrounding the adsorbent. More specifically, the selectivity can be defined as expressed in equation 4 below,
wherein B is the selectivity, x and y are the components whose selectivities are being compared, the quantity (x/y), represents the ratio of the volumetric concentrations of the components at and y which are adsorbed within the adsorbent, the quantity (x/y represents the ratio of the volumetric concentrations of the components x and y which are present in the external phase surrounding the adsorbent. The selectivity of the adsorbent is measured at what is considered to be equilibrium conditions which require that there be no net transfer of materials between the adsorbed and unadsorbed phases when the selectivity is measured. The equilibrium conditions are most easily achieved by passage of a known composition containing the two components, which are to be compared in selectivities, through a bed of the selected zeolitic adsorbent and continuing passage of the feed stream through the zeolitic adsorbed bed until the composition of the material passing out of the adsorbent bed is substantially identical to the composition of feed material passing into the adsorbent bed; Both the feed'and effluent streams having the same composition, a state of equilibrium conditions is achieved and there is no net flow between the adsorbed and external phases of the two components. After equilibrium conditions have been reached, the feed or external stream is drained off from the adsorbent'bed and a desorbent material is then passed through the adsorbent bed to displace essentially all the adsorbed material present within the crystalline aluminosilicate adsorbent. By employing suitable chromatographic equipment to monitor the effluent stream composition and knowing the rates of the material passing into the zeolitic adsorbent bed, it is possible to determine the composition and quantity of the material that was adsorbed by the adsorbent. Upon determination of the composition of the adsorbed material present within the adsorbent and knowing the composition of the feed stream when equilibrium conditions were achieved, it is possible to determine theselectivity as defined in equation 4.
In referring to equation 4 above, a selectivity greater than unity is an indication of selective adsorption by the adsorbent of component x while a selectivity less than unity would indicate a preference of the adsorbent for adsorbing component y. Selectivities. approaching unity indicate that there is no preference by the adsorbent being tested for either of the two components being measured for relative selectivity.
In separating the olefinic and aromatic components present in the effluent stream from the dehydrogenation reaction zone, both liquid and vapor phase operations may be employed with slight differences in the ability to separate the aromatic and olefinic hydrocarbons when operating at either liquid or vapor operations. Liquid phase operations in general require lower temperatures which, when olefinic hydrocarbons are present, appear to be advantageous in that the possibility of polymerization or side reactions occurring during separation is reduced. Therefore, it is preferred, but not necessarily required, that in operations of the adsorptiveseparation zone of the combination process of this invention that liquid phase operations should be effected. In commercial operations where efficiency of operations is of importance, swing-bed or simulated moving bed countercurrent operational techniques may be used to provide continuous paraffin raftinate and extract olefin streams. In general, the swing-bed type operations are effected by employing a number of individual chambers containing the preferred crystalline aluminosilicate adsorbent and, through a predetermined programmed operatiomfeeding dehydrogenation zone effluent into preselected chambers to effect adsorption of the olefinic and aromatic materials produced in the dehydrogenation zone. Simultaneously, a number of preselected adsorbent chambers are contacted with a desorbent material to effectively displace the olefinic hydrocarbons present within the zeolitic adsorbent. By alternating feed and desorbent streams to individual adsorbent chambers, a continuous production of a concentrated olefinic hydrocarbon stream is effected as is also the continuous production of a paraffinic raffinate material which is then recycled to the dehydrogenation zone.
The adsorptive-separation zone can be operated in both vapor or liquid phase with the liquid phase generally preferred. Temperatures of operation can include the range of from 25 C. to about 250 C. and pressures of operation can be within the range of from about atmospheric to about l,000 p.s.1.g.
The dehydrogenation zone operating conditions can include a temperature within the range of from about 760 F. to about l,l0O F a pressure within the range of from about at mospheric to about p.s.i.g., a liquid hourly space velocity within the range of from about 1.0 to 40.0
BRIEF DESCRIPTION OF THE DRAWING Referring to the attached FIGS. a description of the process of this invention is given.
FIG. 1 shows the overall flow scheme of the dehydrogenation zone and adsorptive-separation zone combination while FIG. 2 illustrates in greater detail the flow in the adsorptive separation zone when employing continuously a simulated countercurrent flow adsorptive-separation process to separate olefins and aromatics from paraffinic material.
Referring to FIG. 1, the normal parafiinic feed stream passes through line 1 in admixture with recycle paraffinic material flowing through line 10 and recycle hydrogen material flowing through line 3 and passes into the dehydrogenation zone 11. A portion of the feed material passing through line I into zone 11 is dehydrogenated and leaves the dehydrogenation zone via line 2 and passes into separation zone 12. Separation zone 12 essentially separates gaseous and liquid materials allowing any excess hydrogen produced in the dehydrogenation zone to be removed from the process via line 4. A pressure control means may be connected to line 4 to maintain a steady dehydrogenation zone pressure and to eliminate pressure buildups because of the production of hydrogen during the dehydrogenation reaction. The liquid material separated from the gaseous material flowing out of the dehydrogenation zone passes through line 5 into a midstage fractionation zone 13. The function of the fractionation zone 13 is to separate any light hydrocarbons produced in the dehydrogenation zone and not vented as gaseous hydrocarbons via line 4 in the separation zone 12. The light hydrocarbons are separated from the heavier materials passing into the fractionation zone 13 and pass through line 6 as an overhead material. The heavy materials which consist of the materials boiling at a temperature above the temperature of the normal feed paraffins are separated from the majority of the material passing into fractionation zone via line 7 and are withdrawn as a bottoms stream. Substantially all the feed to fractionation zone 13 is removed therefrom via line 8 as a sidecut fraction and passes through line 8 into adsorptive-separation zone 14. Adsorptive separation zone I4 as previously mentioned can comprise swing-bed type operations or continuous countercurrent flow bed operations in which olefinic and aromatic materials produced in the dehydrogenation zone ll are separated from the unreacted parafiinic materials. The unreacted parafflnic hydrocarbons which pass through the dehydrogenation zone 11 and are separated from the aromatic and olefinic materials in the adsorptive separation zone 14 are returned to the dehydrogenation zone via paraffin recycle stream 10. The extract stream which comprises substantially all mono-olefinic hydrocarbons and a small portion of side reaction produced aromatic hydrocarbons is removed from the adso'rptive-separation zone via line 9 and can be recovered as a product olefin material or can be passed to a reaction zone to be subsequently reacted in a process requiring high purity monmolefinic hydrocarbons.
FIG. 2 represents a detailed process flow for a continuous simulated moving bed countercurrent adsorption-separation process.
Referring to the attached FIG. 2, a detailed explanation of a continuous simulated moving bed separation zone I4 is given. Lines 28, 8, 29 and 15 are connected to flow directing valve 16 and have flow-controlling valves 30, 31, 32 and 33 attached for independent control of the individual raffinate, feed, extract and desorbent flow rates. Line 8 carries the feed from fractionation zone 13 of FIG. 1 to the flow director 16 and subsequently to adsorption column 17.
Line 28 of the attached FIG. 2 carries the relatively less sorbed components of the feed which comprise the paraffinic hydrocarbons. The less selectively sorbed components of the feed flow through line 18 from column 17 at a rate which is controlled by valve 30. The raffinate material flowing from column 17, in addition to paraffinic hydrocarbons, contains desorbent material which was displaced from the sorbent by the normal olefins in the feed passed into the adsorption column via line 8. The raffinate material flowing from column 17 is separated by an external fractionation zone to yield desorbent and normal paraffin fractions; the desorbent is recycled to line 15 for reuse and the paraffinic hydrocarbon is collected and recycled to the dehydrogenation zone via line of FIG. 1.
Line 29 of the attached FIG. 2 carries extract material from column 17 at a rate controlled by valve 32. The extract material comprises feed-olefins, aromatics and desorbent material and is the resultant stream formed by displacement of adsorbed feed olefins and aromatics by the desorbent stream flowing through line 15. The extract stream flowing through line 29 is separated into an olefin product stream and a desorbent stream in an external fractionation zone. The feed olefin is recovered as product from line 9 of FIG. 1 and desorbent material is preferably recycled to line for reuse.
Flow directing valve 16 of FIG. 2 connects lines 28, 8, 29 and 15 to lines 18, 19, 20, 21, 22, 23, 24 and 25 which are connected to column 17. Lines 18 through 25 enter the column 17 through inlet-outlet ports that are located between the eight individual fixed beds in a preferably narrow portion of the column 17. For example, line 23 enters the column 17 through port 27 between beds 5 and 6. Flow directing valve 16 can comprise a multivalve manifold arrangement, a rotary multiport valve or any other suitable flow directing mechanism that will, in a programmed manner, direct flow of the feed (line 8) and desorbent (line 15) streams into the column and the raffinate (line 28) and extract (line 29) streams out ofthe column.
In FIG. 2, the feed flows through line 8 to flow director 16 wherein the feed is sent through line to column 17; the desorbent flows through line 15 to the flow director which sends the desorbent through line 24 to the column; the raffinate stream flows from column 17 through lines 34 and 18 to the flow director wherein the raffinate stream is sent through line 28 to external fractionation facilities; the extract stream flows from column 17 through line 22 to the flow director which sends the extract stream through line 29 to external fractionation facilities. The streams flowing into and out of the process as described above comprise a single cycle (Cycle 1 of table I) of an operation which will vary in length of time according to the feed composition, required product purity, sorbent properties, etc.
Table I below indicates the locations of the feed, raffinate, extract and desorbent streams during the individual cycles used in the continuous operation of column 17.
Lines through which material is flowing from flow directing valve to absorption column (refer to FIGURE 2) Cycle 8 is the last cycle in completing one sequence of operations. After cycle 8 is completed, cycle I is commenced.
NOTE: R Raflinate stream; E Extract stream; D Desorbent stream; F=Feed stream.
As can be seen in table I, the raffinate, feed, extract and desorbent streams that flow to column 17 are advanced equally in the same direction when advancing to the next cycle of operations. It should be understood that any number of cycles greater than four may be used and that the number of cycles required for one complete sequence of operations depends on the number of individual inlet-outlet ports that the column contains.
As can be seen in table I, there are only four of the total of eight lines entering the column 6 that are in use during a given cycle. For example, in cycle I of table I, lines 18, 20, 22 and 24 are in use while lines 19, 21, 23 and 25 are not being used. The flow distributor 16 is constructed in a manner so that the lines not having material flowing through them during a given cycle, i.e. lines 19, 21, 23 and 25 of cycle I, may be plugged or blocked off at either or both the flow director or column ends thereby stopping flow through these lines. Alternately purge streams of feed, extract, raffinate, desorbent or other material may be passed through the lines which are not in use during a particular cycle to prevent contamination of the various streams with material present in one of the lines 18 through 25 from a previous cycle of operations.
The adsorption column 17 of FIG. 2 is a plurality of serially connected fixed beds containing a selected solid adsorbent that has a higher sorbing affinity for olefins and aromatics than for corresponding paraffins of the same carbon range. Column 17 contains eight fixed beds numbered through I through 8 with the terminal beds (beds 2' and 3) connected by line 34. Pump 26 in line 34 provides a means for circulating liquid or vapor from the top of column 17 to the bottom thereof. The pump-around system gives the fluid in column 6 a unidirectional flow which relative to the stationary solid sorbent in the eight beds of column 17 flows from bed 3' to bed 2 via beds 4, 5, 6', 7', 8 and 1. Relative to bed 3, bed 4' is in a downstream direction; relative to bed 4', bed 3 is in an upstream direction by virtue of the direction of fluid flowing through the separate beds.
To reduce the operation of the adsorption column to relatively simplified terms, the column can be thought to be operating in continuous counter flow of liquid and said adsorbent with the overall separation of olefins and paraffins being effected by four separate zones.
Zone I is the series of beds located between the port of feed introduction downstream to the port of raffinate withdrawal; zone II is the series of beds located between the port of extract withdrawal downstream to the port of feed introduction; zone III is the series of beds located between the port of desorbent introduction downstream to the port of extract withdrawal; and, zone IV is the series of beds located between the port of raffinate withdrawal downstream to the port of desorbent introduction. As mentioned previously, the ports of feed and desorbent introduction and the ports of raffinate and extract withdrawal are advanced equally and essentially simultaneously in a downstream direction during the various cycles of operation shown in table I. Consequently, zones I, II, III and IV are advanced equally and simultaneously in a downstream direction as the inlet and outlet ports are so advanced.
Table II indicates the location of the individual zones in the series of beds in the adsorption column for the individual cycles used in the continuous operation of the adsorption column.
TABLE II.-ZONE POSITIONS IN THE ABSORPTION COLUMN FOR VARIOUS CYCLES OF OPERATION Cycle Adsorbent bed h in column 1 2 3 I 5 6 7 8 II II III III IV IV I I II II III III IV IV I I II II III III IV IV I I II II III III IV I\ I I II II III III I\' IV I I II II III III I,\' IV I I II II III III IV I\ I I Cycles 1-8 of Table I.
In zone I, olefins and side reaction aromatics in the feed are adsorbed by the solid adsorbent displacing the previously adsorbed desorbent. During the normal course of operation, zone I is shifted as previously described. In using zone I as a reference point, when zone I shifts to the next bed position (see table ll) the solids that leave zone I enter zone [1. The adsorbent that is entering zone ll carries olefins and aromatics adsorbed from the feed and other hydrocarbons comprising paraft'ins from the feed. in zone ll, the paraffins from the feed and most other nonolefinic and nonaromatic hydrocarbons are displaced from the solid by desorbent. Any olefinic or aromatic hydrocarbons from the feed that are desorbed from the adsorbent in zone ii are readsorbed in zone I. The adsorbent in zone [I] carries primarily olefins and aromatics from the feed and some desorbent is contacted with a large excess of desor bent which displaces essentially all of the olefins and aromatics from the feed that were carried on the adsorbent. When zone lll shifts to its new location in the adsorption column. the adsorbent that leaves this zone and enters zone IV carries primarily desorbent which can be made available for reuse in zone ill by contacting the adsorbent with a portion of the raffinate. in zone IV, the displacing of most of the desorbent by raffinate material is accomplished. The rafi'inate flow rate into zone W from zone I is controlled so that the raffinate material flowing into zone IV is completely adsorbed and does not pass entirely through zone IV.
In starting the process (cycle 1 of tables I and II), a feed stock such as the dehydrogenation zone effluent mixture is charged into the process flow through line 8, at a rate controlled by valve 31, through flow director 16 and into line 20 which carries the feed into the column at the port located at the preferably narrow portion of the column located between beds 1' and 8'. The feed entering the column through the port of line 20 flows in a downstream direction into bed 1 wherein the olefins and aromatics in the feed and some paraffms are adsorbed by the solid adsorbent. Simultaneously, part of the desorbent present in the solid-adsorbent pores from a previous cycle of operation is displaced from the adsorbent. The less strongly adsorbed paraffins occupy the interstitial void spaces between the solid particles of the adsorbent and eventually flow downstream towards bed 2' and to line 34 which allows a portion of the raffinate stream (paraffin and desorbent mixture) to be withdrawn from the column via line 18, the flow director 16 and line 28, the remaining raffinate flows through pump around line 34 to bed 3 of zone IV. The solid adsorbent in bed 1' which contains, in addition to adsorbed normal olefins and aromatics, a quantity of paraffins which can be displaced from the adsorbent by desorbent, which from a prior cycle, is contained in upstream beds 7 and 8. The paraffins which are displaced from the solid adsorbent in bed 1' by desorbent material flow in a downstreamdirection towards bed 2. Unavoidably, some of the olefins and aromatics adsorbed on the solid adsorbent in bed 1 are displaced at the same time. The flow rate of liquid flowing into bed 1' from bed 8 can be adjusted to displace substantially all of the parafiins adsorbed by the adsorbent in beds 1 and 2', without simultaneously washing out all of the more tenaciously adsorbedolefins and aromatic hydrocarbons. Any olefins and aromatics which are desorbed in bed 1' are readsorbed in bed 2.
The normal paraffin hydrocarbons together with desorbent material are the principal materials withdrawn from bed 1', assis s b t 29n ie 19 n rinshs. i lw y olefins and aromatics in the material entering bed 2' can be adsorbed. The stream passing out of bed 2' through line 34 comprises principally the paraffins and desorbent. A portion of the fluid effluent from bed 2' in line 34 passes through line 18 to flow director 16 and through the rafi'mate line 28, the raffinate flow rate out of the column is controlled by valve 30 in line 28. The remaining portion of effluent from bed 2 flows through line 34 into bed 3. Line 34 connects the terminal beds 2' and 3' and allows continuous unidirectional flow of liquid through the column.
The solid adsorbent in beds 3' and 4' contains within its pores essentially only adsorbed desorbent which is present and desorbent material. The paraffins are adsorbed on the adsorbent in bed 3' displacing desorbent material downstream to bed 4 and bed 5'. The flow rate of the raffinate material into bed 3' is adjusted so that paraffins are completely adsorbed on the adsorbent before reaching the outlet of bed 4'. Otherwise,
the parafi'ins would contaminate the olefinic product in the extract stream.
The solid adsorbent in beds 5 and 6 contains adsorbed olefins, aromatics and desorbent from a previous cycle of operations. The aromatics and olefins which have been selectively adsorbed from the feed are displaced by desorbent material flowing through line 15, at a rate controlled by valve 33, to the flow distributor, through line 24 and out of the port of line 24 between beds 4' and 5'. The desorbent upon entering the column flows in a downstream direction into beds 5 and 6' displacing the oleflnic product. The desorbent and olefinic material and the small quantity of aromatic side reaction products which comprise the extract stream flow out of column 17 at the port between beds 6 and 7' through line 22 to the flow distributor l6 and through line 29 at a rate controlled by valve 32 in line 29. A portion of the extract material flows past the port of line 22 into the next downstream bed 7'. Any olefinic product passing into bed 7' is adsorbed by the solid adsorbent inbed 7'. The desorbent material flowing through bed 7' into bed 8 tends to flush any paraffins that are carried by the solid adsorbent when the feed line is shifted from line 20 (cycle 1 to line [9 (cycle 2).
Generally, the parafiinic hydrocarbon portion of the rafflnate stream not withdrawn from column 17 through the raffinate withdrawal line does not contaminate the stream of liquid flowing beyond the first downstream bed from the port of raffinate withdrawal. The same conditions apply for the olefin portion of the extract stream.
The above-described flow of feed and desorbent streams into the column and extract and rafflnate streams out of the column comprise cycle 1 of table l. Cycle 2 of table I is then executed with the feed line switching from line 20 to 19, the rafflnate line switching from line 18 to 25, the desorbent line switching from line 24 to 23, and the extract line switching from line 22 to 21. The feed, raffinate, desorbent and extract lines are advanced in the direction of net liquid flow through the column with the valves 31, 30, 33 and 32 altering the input and output flow rates to achieve desired extract and raffinate purities.
Lines l8, 19, 20, 21, 22, 23, 24 and 25 carry different streams to and from the adsorption column during the individual cycles of operation of the flow director. For example, line 22 during cycle l of table I carries the extract stream, while during cycle 5, line 22 carries raffinate material. To eliminate the contamination of raffinate and extract streams, it is understood that a method of flushing the lines 18 through 25 is preferred. A preferred method of flushing lines 18 through 25 is to pump desorbent material through the line immediately upstream of the feed inlet into the adsorption column as described in U.S. Pat. No. 3,201,491. in flushing the line immediately upstream from the feed line, the extract which eventually will be flowing out of adsorption column through a previously flushed line will not be contaminated with the feed stock components not desired in the extract. This increases product (extract) purity and favorably affects the quality of the extract material.
EXAMPLE i This example illustrates the principal benefits of using the combination process of this invention to effectively reduce the buildup of aromatic hydrocarbons in the paraffmic recycle stream passing into the dehydrogenation zone thereby eliminating catalyst deactivation associated with the presence of aromatic hydrocarbons in the dehydrogenation reaction zone feed stock.
The catalyst used in the dehydrogenation reactor was prepared according to the teachings of U.S. Pat. No. 3,291,755 and contained 0.76 wt. percent Pt., 0.041 wt. per
cent As, 0.55 wt. percent Li, all composited on a gamma alumina carrier.
The dehydrogenation reactor was operated at a liquid hourly space velocity based on combined feed (fresh feed recycle paraffins) of about 28.0, a pressure at the outlet of the reactor of 30 p.s.i.g., a gaseous hydrogen to hydrocarbon mole ratio in the reactor of 9:1 and a temperature of from about 850 F., to about 950 F in order to maintain a 10 wt. percent conversion of combined feed to essentially mono-olefinic hydrocarbons.
In order to demonstrate the effects of aromatic hydrocarbons on the catalyst when present in a feed stream passing into the dehydrogenation reaction zone, two separate tests were conducted. ln Test A, a combined feed ratio, (CFR volume of fresh feed recycle paraffins/fresh feed) of 10.0 was maintained. A recycle paraffin stream was used which resembled that which would be obtained when employing the adsorptiveseparation unit used in the combination process of this invention to separate olefins and most of the side-reaction produced aromatics from the unreacted paraffins present in the dehydrogenation zone effluent. ln test B, a recycle paraffin stream was employed which contained essentially all of the aromatic hydrocarbons produced by the dehydrogenation zone.
The fresh feed charge stock used in tests A and B had a APl gravity of 55.8", with an initial boiling point of 354 F., a 10 percent boiling point of 400 F., 50 percent boiling point of 413 F., a 90 percent boiling point of 444 F., and an end boiling point of 459'- F. This feed stock comprises primarily C through C normal paraffms with specific composition as follows: 0.3 wt. percent n-C 26.4 wt. percent n-C 31.2 wt. percent n-C 25.3 wt. percent n-C 13.3 wt. percent n-C 0.4 wt. percent nC, 0.21 aromatics (benzenes and naphthalenes), 1.46 wt. percent monocyclic paraffins, 0.68 dicyclic paraffins and 0.75 wt. percent iso-paraffins. The recycle paraffin stream used in Test A was essentially the same as the feed composition except that it contained about 1.1 wt. percent aromatics present as benzenes and naphthalenes. The recycle stream used in Test B was similar in composition to the feed stream except that it contained about 3.0 wt. percent aromatic hydrocarbons.
Results of the two aforementioned tests are shown.
TABLE III Test A Test B Wt. percent aromatics in paraffin recycle stream at 105 b.p.p. catalyst life Deactlvatlon rate at 105 l .p.p.. F./BPP 1 l Barrels of combined feed charged per pound of catalyst. 2 Degrees increase in catalyst temperature per barrel of combined feed charged per pound of catalyst at conversion.
As can be seen from the results of table 111, a reduction in the concentration of aromatic hydrocarbons contacting the dehydrogenation catalyst increased the stability of the dehydrogenation catalyst by decreasing the rate of deactivation from l.6 F./BPP to 0.09" FJBPP. a decrease in the deactivation rate at 105 8?? catalyst life of over 17 fold.
EXAMPLE ll A C through C dehydrogenation reactor effluent was used as a feed in a series of separation experiments used to verify the ability of a selected adsorbent to selectively separate a mixture of olefins for the process of this invention. The dehydrogenation reactor effluent composition is shown in table IV.
n-C olefin n-C paraffin n-C olefin n-C paraflin n-C olefin n-C parafiin n-C olefin C1 paraffin Total normal olefin Total normal parafiln Total non-normals The dehydrogenation reactor effluent was passed through a bed of 320 cc. of a selected type Y adsorbent at a pressure of 300 p.s.i.g., and a liquid hourly space velocity of 2.8. After the adsorbent appeared to be fully loaded with the olefins from the dehydrogenation reactor effluent, a flush stream of isoand normal pentane was passed through the adsorbent bed to flush away paraffins remaining in the interstitial voids between the adsorbent particles. After the paraffms from the dehydrogenation reactor effluent were removed, a desorbent stream was passed through the adsorbent bed to remove the selectively sorbed olefins and aromatics. The recovered mixture which comprised desorbent and the selectively adsorbed olefins and aromatics was fractionated to remove the desorbent material. The remaining olefinic extract material was analyzed using gas-liquid chromatographic methods.
Five individual tests were run using the feedstock of table IV and the same general procedures described above. The operating conditions for the individual tests are summarized below:
Test C: A sodium form type Y zeolite adsorbent which was calcined for 2 hours at 450 C, was used. The adsorption and desorption operations were carried out at 300 p.s.i.g., and 150 C. The desorbent material which was used to displace the adsorbed olefins and aromatics was essentially percent octane- 1.
Test D: The adsorbent employed in this test was a silverexchanged type Y structured zeolite which after drying at 450 C, for 2 hours, contained about 9.85 wt. percent silver, as the element. Adsorption and desorption operations were effected at 300 p.s.i.g., and 100 C., and the desorbent used comprised 16 vol. percent octane-l in isooctane.
Test E: Same as test 2 except a 100 vol. percent octane-l desorbent was used.
Tests F and G: The adsorbent was a silver-exchanged type Y structured zeolite which contained about 8.5 wt. percent silver after drying for 2 hours at 450 C. Adsorption and desorption conditions were effected at 300 p.s.i.g., and 100 C. A two-step desorption was employed using a first desorbent of about 100 vol. percent octane-l followed by a second desorbent of 1.5 vol. percent octane-1T1; lso-octane. The results of the above five tests are shown in table Vfollowing.
TABLE V.OLEFINIC EXTRACT ANALYSIS Test C I) E F G Total vol. of feed passed through adsorbent bed. cc 2,595 2,270 2,439 1,977 1,970 Purity of olefinie material recovered 1 vol. percent 17.7 96.8 05.0 97.8 98.4 GLC analysis, vol. percen C Trace 3.0 0.1 0.1 Trace n-Cm paraffin ..Tracc 0.2 0.1
.i I... mono-olefin 2. 8 1.3 5.6 1. 6 *DESCRIPTION OF THE fiiziriiniapiiiiaofiiiinms ll-Cupalam11 x.- 4.6 0.9 0.8 0.2 0.3 l
ii-Cii mono-olefins. 0.5 23. 7 17. 5 23. 4 23. 3
f gggggf 8:8 5:: 3g 8-; In a broad embodiment, the invention comprises a process ilCia mono-olefins. 7.1 25. 2 27. 4 23. 25.1 5 for the production and recovery of normal olefins which figgfii g 9 2:3 81% 1:3 9:5? 3:2 process comprises the steps of: (a) contacting a normal parafn-C paraflins.. m nu fin-containing hydrocarbon feed stream with a dehydrogenation catalyst in a dehydrogenation zone at dehydrogenation MS analysis of 015+, v lpercont conditions to effect production of a major reaction product of 950 normal mono-olefins and additionally producing, as a minor Olefins andparamns Aromatlc types: side-reaction product, aromatic type hydrocarbons; (b)
B 1 1b 4.4 9.9 15355358; 0,5 3,3 withdrawing an effluent stream comprising olefins, paraffins, rii fiififaifiiiifiiffi f935* ???ii:iititj'. $523 ii mm 5*, hydmgenafiflfl W separating said effluent into a liquid phase comprising olefins, Total awmatics 15 paraffins and aromatics and a gaseous phase comprising f 1 Based n (gm flimilglh C1 extract material recovered after separation hydrogen; (d) contacting at least a portion of said liquid phase rom 6801 en 1119. 61 a NOTE. Tests D and G were the only two tests in which the (315+ with a crystalline aluminosilicate adsorbent, having uniform terial was analyzed. pore openings of from about 6 to about 15 angstrom units and containing exchangeable cations selected from the group con- XA PLE lll E M sisting of Group IA, Group "A, Group IB and Group "B In this example, the dehydrogenation reaction zone and the metals, at adsorption conditions to effect the adsorption of adsorptive-separation zone are operated in an integrated said normal mono-oiefins and aromatic hydrocarbons within manner as indicated in FIG. 1. The dehydrogenation zone said adsorbent; (e) withdrawing less selectively adsorbed rafcontains a catalyst similar to the one employed in example I finate material comprising saturated components of said liquid and is operated at conditions which include a p.s.i.g., reacphase effluent from said adsorbent; (f) recycling at least a portor outlet pressure, a mole ratio of gaseous hydrogen to the tion of said raffinate material into said dehydrogenation zone; hydrocarbon present within the reaction zone of about 8:1, an (g) contacting said adsorbent with a desorbent material at inlet reactor temperature of about 880 F., a combined feed desorption conditions to effect displacement of said adsorbed ratio (fresh feed recycle paraffin/fresh feed) of about 1 1.0, a 30 olefins and aromatic hydrocarbons; and (h) withdrawing from liquid hourly space velocity based on the combined feed of said adsorbent an extract stream comprising desorbent materiabout.28.0 and 2,000 wt. p. .m. water in the combined feed. al. a omatic hydrocarbons and olefins. The dehydrogenation zone effluent is passed into the adsorpl l im s my in nti n: tive-separation zone which contains a type Y zeolite which has A P B the P ion an recovery of normal been essentially totally exchanged with a water soluble potasol fins hich process comprises the steps of: sium salt. The potassium type Y zeolite adsorbs both olefinic a contacting a normal paraffin-containing hydrocarbon and aromatic components of the dehydrogenation zone efd tream with a dehydrogenation catalyst in a fluent and is essentially inert towards olefin polymerization. dehydrogenation zone at dehydrogenation conditions to The aromatic materials adsorbed by the adsorbent are effect production of a major reaction product of normal desorbed from the sieve by employing a desorbent that comm n l fin n i i y pr ing. as a m n r s eprises at least 50 volume percent of linear olefinic material. reaction P afomaiic hy r The adsorptive-separation unit is operated at about 25 p.s.i.g., h r g n effluenl stream p ng olefins. p fwith an operating temperature of about 275 F. fins, aromatics and hydrogen from said dehydrogenation The feed material passing into the dehydrogenation zone is Zone; essentiallya kerosene-boiling range material boiling in the C separating said effluent into a liquid phase comprising to (I carbon number range, olefins, paraffins and aromatics and a gaseous phase com- The combination process is started up and lined out prior to p i g hy g n; determining the individual stream compositions. Table V! nt ting at l ast a p rtion of said liquid phase with a shows the various input, output and internal stream rates for y li aluminosilicaie adsorbent, having n rm the combination ro e hown in FIG, 1, pore openings of from about 6 to about 15 angstrom units TABLE VI Line 1. Line 10. Line 4. Line 6,
fresh recycle off over- Line 7, Line 9, Pounds/day feed paraflin gas head bottoms extrac Component:
H, 210 C 5 C2 r 5 C: 4 C4 a 2 C5 1 Cs-Cn Nonial paraflins:
l0 v A a 011- 5, 715 67, 987 Cu. 5,590 59,278 013-. 3,973 37,031 g...- .l. 1,941 12,982
15 s c v Iso-i-cycloparaflins 193 2, 337 Aron:iat:ics. ....v 88 427 Normal olefins:
CXL 564 C11- 556 C13- 394 CH- 153v Iso+cycloparaffins 19 Di-oleflns. 63 .4
Total, lb./day 17,500 181, 791
The above examples are offered as being illustrative of the h and containing exchangeable cations selected from the combination process this invention and are not to be congroup consisting of Group IA, Group "A, Group IB and strued as limiting the scope thereof. Group IlB metals, at adsorption conditions to effect the adsorption of said normal mono-olefins and aromatic hydrocarbonswithin said adsorbent;
e withdrawing less selectively adsorbed raffinate material comprising saturated components of said liquid phase effluent from said adsorbent;
f. recycling at least a portion of said raffinate material into said dehydrogenation zone.
g. contacting said adsorbent with a desorbent material at desorption conditions to effect displacement of said adsorbed olefins and aromatic hydrocarbons; and,
h. withdrawing from said adsorbent an extract stream comprising desorbant material, aromatic hydrocarbons and olefins.
2. The process of claim 1 further characterized in that said dehydrogenation conditions include a temperature included within the range of from about 750 F. to about l,000 F. and a pressure included within the range of from about atmospheric to about 200 p.s.i.g.
3. The process of claim 1 further characterized in that said desorbent material contains normal olefinic hydrocarbons in a concentration of from about 25 to about l liquid volume percent of said desorbent material.
4. The process of claim 1 further characterized in that said dehydrogenation catalyst comprises an alumina base containing from about 0.10 to about 5.0 wt. percent of a Group Vlll metal and from about 0.01 to about 5.0 wt. percent of a metal selected from the group consisting of Group IA, Group A metals, rhenium and germanium metals.
5. The process of claim 4 further characterized in that said dehydrogenation catalyst comprises a lithium modified alumina base having from about 0.10 to about 2.5 wt. percent platinum and from about 0.01 to about 2.5 wt. percent lithium composited thereon.
6. The process of claim 1 further characterized in that said normal paraffin-containing hydrocarbon feed contains paraffins having from about 6 to about 20 carbon atoms per molecule.
7. The process of claim 6 further characterized in that hydrocarbon feed contains paraffms having from about 10 to about carbon atoms per molecule.
8. The process of claim 1 further characterized in that said adsorption and desorption conditions include a temperature included within the range of from about 70 F. to about 350 F. and a pressure included within the range of from about atmospheric to about 500 p.s.i.g.
9. The process of claim 8 further characterized in that said adsorption and desorption conditions include liquid phase operation.
10. The process of claim 1 further characterized in that said crystalline aluminosilicate zeolite IS selected from the group consisting of type X and type Y structured zeolites.
II. The process of claim [0 further characterized in that said crystalline aluminosilicate comprises type X structured zeolite.
12. The process of claim 11 further characterized in that said crystalline aluminosilicate contains exchangeable cations selected from the group consisting of copper, potassium and sodium.
13. The process of claim 10 further characterized in that said crystalline aluminosilicate comprises type Y structured zeolite.
14. The process of claim 13 further characterized in that said crystalline aluminosilicate contains exchangeable cations selected from the group consisting of silver, copper, potassium and sodium.
15 A process for the production and recovery of normal olefins which process comprises the steps of:
a. contacting a normal paraffin-containing hydrocarbon feed stream with a dehydrogenation catalyst comprising an alumina base containing from about 0.10 to about 5.0 wt. percent of a Group Vlll metal and of from about 0.01 to about 5.0 wt. percent of a metal selected from the group consisting of Group IA and Group "A metals, in a dehydrogenation zone at dehydrogenation conditions to effect production of a major reaction product of normal mono-olefins and additionally producing, as a minor side reaction product, aromatic hydrocarbons;
b. withdrawing an effluent stream comprising unreacted normal paraffins, mono-olefinic hydrocarbons, side reaction product aromatic hydrocarbons and hydrogen from said dehydrogenation zone;
c. separating said effluent into a liquid phase comprising olefins, paraffins, and aromatics and a gaseous phase comprising hydrogen;
. contacting at least a portion of said liquid effluent with a crystalline aluminosilicate adsorbent, selected from the group consisting of type X and type Y structured zeolites and containing exchangeable cations selected from the group consisting of Group lA, Group "A, Group [B and Group [18 metals, at adsorption conditions to effect the adsorption of said normal mono-olefins and aromatic hydrocarbons within said adsorbent while displacing a portion of desorbent materials present within the adsorbent from a previous desorption step;
e. withdrawing from said adsorbent a stream comprising desorbent material and parafflnic hydrocarbons;
f. contacting said adsorbent with a desorbent material at desorption conditions to effect displacement of said adsorbed olefms and aromatic hydrocarbons from said adsorbent;
g. withdrawing from said adsorbent an extract stream comprising desorbent material, aromatic hydrocarbons and olefins;
h. separating said desorbent from the mixture of paraffins and-desorbent and recycling at least a portion of said paraffins into said dehydrogenation zone.
16. The process of claim 15 further characterized in that said dehydrogenation catalyst comprises a lithium modified alumina base having from about 0.10 to about 2.5 wt. percent platinum and from about 0.01 to about 2.5 wt. percent lithium composited thereon.
17. The process of claim 15 further characterized in that said dehydrogenation conditions include a temperature included within the range of from about 750 F. to about l,000 F. and a pressure included within the range of from about atmospheric to about 200 p.s.i.g.
18. The process of claim 15 further characterized in that said normal paraffin-containing hydrocarbon feed contains paraffins having from about 6 to about 20 carbon atoms per molecule.
19. The process of claim 18 further characterized in that said hydrocarbon feed contains paraffins having from about 10 to about 15 carbon atoms per molecule.
20. The process of claim 15 further characterized in that said adsorption and desorption conditions include a temperature included within the range of from about 70 F. to about 400 F. and a pressure included within the range offrom about atmospheric to about 500 p.s.i.g.
21. The process of claim 20 further characterized in that said adsorption and desorption conditions include liquid phase operations.
22. The process of claim 15 further characterized in that said desorbent material contains normal olefinic hydrocarbons in a concentration of from about 25 to about l00 liquid vol. percent of said desorbent material.
23. The process of claim 20 further characterized in that said normal olefinic hydrocarbons present in said desorbent material have from about 4 to about 15 carbon atoms per molecule.
24. The process of claim 15 further characterized in that said crystalline aluminosilicate comprises type X structured zeolites.
25. The process of claim 24 further characterized in that said crystalline aluminosilicate contains exchangeable cations 19 20 selected from the group consisting of silver. copper. potassium 27. The process of claim 26 further characterized in that and sodium. said crystalline aluminosilicate contains exchangeable cations 26. The process of claim further characterized in that selected from the group consisting of silver. copper. potassium said crystalline aluminosilicate comprises type Y structured 5 and iumzeolites. ne e e