|Publication number||US3496095 A|
|Publication date||Feb 17, 1970|
|Filing date||Mar 4, 1968|
|Priority date||Mar 4, 1968|
|Also published as||CA926800A, CA926800A1, DE1909840A1|
|Publication number||US 3496095 A, US 3496095A, US-A-3496095, US3496095 A, US3496095A|
|Inventors||Emil H Lewis|
|Original Assignee||Exxon Research Engineering Co|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (6), Referenced by (11), Classifications (13)|
|External Links: USPTO, USPTO Assignment, Espacenet|
Feb. 17, 1970 Y E. H. LEwls PRooEss Foa UPGRADING STEAM GRACKED FRACTIoNs 5 Sheets-*Sheet l.
Filed March 4. 1968 INVENTOR. EMIL H. LEWIS ATTORNEY r wm Q59.. E
s 3 d v w l o H wzim m25 5E...
E. H. LEWIS Feb. 17, 1970 PROCESS FOR UPGRADING STEAM CRACKED FRACTIONS Filed March 4, 1968 3 Sheets-Sheet 2 mzmNzmm E nu Feb. 17, 1970 E. H. LEWIS 3,496,095
PROCESS FOR UPGRADING STEAM CRACKED lFRACTIONS Filed March 4. 1968 3 Sheets-Sheet 5 LLI E l o Il) 3 m m EMIL H. LEWIS AT TORNE Y United States Patent C) U.S. Cl. 208-57 10 Claims ABSTRACT OF THE DISCLOSURE This disclosure relates to an integrated process for refining steam cracked fractions featuring two stage hydrogenation to maximize the production of motor fuel components, a BTX fraction and/or benzene. The first stage catalyst comprising nickel and tungsten is a critical feature of the invention.
In recent years, the steam cracking process has become increasingly popular as a source of ethylene and other olefin gases. Feedstocks ranging from light hydrocarbon gases to heavy oils such as gas oils and crude oils are passed with steam through a high temperature pyrolysis furnace where cracking takes place. The steam cracking process is designed and operated to produce a maximum amount of ethylene and propylene. However, liquid products including a raW steam cracked naphtha or pyrolysis gasoline are produced in varying amounts depending on feed characteristics and operating conditions. Typical commercial processes are described in Hydrocarbon Processing, November 1965, vol. 44, No. 11, pp. 203-208.
Steam cracking is also employed to produce acetylene or a mixture of acetylene and ethylene. This process also produces a quantity of steam cracked naphtha. Commercial processes of this type are described in the previously cited publication at pages 165 and 166.
When the feed to the steam cracking process is a liquid petroleum fraction and particularly when the fraction is a heavy oil such as gas oil or crude oil, the cracked liquid fractions are extremely unstable. Refiners have been using heavier feeds because of the market demand for such steam cracking co-products as propylene, stable steam cracker gasoline and aromatics such as benzene, toluene and xylene.
The liquid effluent from t-he steam cracking of heavy oils is contaminated with olens, diolefins, gum and sulfur compounds. The contaminants are usually removed by a number of treating steps such as redistillation, heat soaking, clay treating, chemical treating, hydrofining and solvent extraction. Each of these steps is costly and reduces the economic incentive for producing motor gasoline components and aromatic hydrocarbons.
The object of this invention is to provide a process for the upgrading of unstable steam cracked fractions which requires a minimum of treating steps and which can be operated to maximize the production of gasoline, a benzene-toluene-xylene (BTX) fraction or benzene alone. Another object of the invention is to provide a process which includes two stages of hydrogenation employing low cost catalysts which have a long onstream life so that continuous operations can be carried out over many months.
Briefly summarizing a raw unstable steam cracked fraction is hydrogenated without pretreatment in a first hydrogenation step in the presence of a catalyst comprising nickel and tungsten on a suitable support, then all or a portion of the partially hydrogenated liquid fraction is hydrogenated in a second hydrogenation step at conditions more severe than those employed in the first step in the presence of a conventional hydroiining catalyst such as suliided cobalt molybdate. Following the second hydrogenation step, the subsequent steps vary according to the products desired. If the product is motor-gasoline, simple distillation is sufficient. If the prime product is a BTX fraction, solvent extraction and any desired number of distillation steps are used. If the process is being operated to produce a maximum quantity of benzene, dealkylation can be used to remove alkyl side chains from alkyl aromatic hydrocarbons.
Further details of the process will be described below with reference to the drawings in which FIGURE 1 is a flow sheet illustrating an embodiment in which the production of motor gasoline is maximized. FIGURE 2 illustrates an embodiment in which the process is operated t0 maximize the production of a BTX fraction.
The process feed is a highly unstable liquid co-product fraction recovered from the steam cracking of a liquid petroleum oil, such as gas oil. The feed contains such materials as cyclopentadiene, styrene, dicyclopentadiene and acetyleneic hydrocarbons. The preferred feed boils in the range of from about 50-500 F., usually in the range of 100-430 F. The feed has a total olefin content of 5-40 wt. percent, including a monoolefin content of 2-20 Wt. percent, a diolefin content of 1-20 wt. percent, an aromatic hydrocarbon content of 40-90 wt. percent, a sulfur content of at least 100* p.p.m. generally ranging from 100 to 4000 p.p.m., a gum content of 1 to 1000 mg./ 100 ml. and a bromine number of at least 50', generally ranging from 50 t0 125. If desired other unstable fractions can be blended with the steam cracked fraction. It is conventional to remove gum and sulfur from such fractions by caustic Wash redistillation and/or heat soaking but it is a feature of this process that such pretreating is eliminated.
Referring to FIGURE l a typical raw steam cracked naphtha is passed by line 1 to hydrotreating reactor 2. The naphtha has the properties shown in column 1 of Table I and thus would be considered very unstable. A gas containing -100% hydrogen is added to the reactor by lines 3 and 1. Any suitable reactor can be used for the first stage hydrogenation step. The reaction is usually started by flowing hot hydrogen-containing gas through the reactor to obtain the desired temperature in the catalyst bed, then adding the steam cracked naphtha iu the liquid phase.
The first stage catalyst is nickel-tungsten on a support as alumina, keiselguhr, silica-free clay, bauxite, mullite, etc. Alumina is the preferred support and the catalyst is sulde. The catalyst can contain from about 2 to about 6 wt. percent nickel; from about 10y to about 20 Wt. percent tungsten and the balance alumina. The tungsten to nickel ratio is preferably about 4 to 1 to 5 to l. The
catalyst can be made from the metals or their salts, e.g., oxides, nitrates, chlorides, etc. Soluble ingredients such as nickel nitrate and ammonium meta or paratungstate are preferred.
It is preferred that the reactor inlet temperature be maintained at less than about 290 F., i.e. 150 to 290 F. With feeds containing significant amounts of diolefins rapid plugging of the catalyst bed will occur at inlet temperatures above about 300 F. The temperature increase across the bed should be limited to about 60 to 400 F. The average reaction temperature should be 225 to 500 F. Reactor pressures of 400 to 1000 p.s.i.g. are suitable. Space velocities range from 0.25 to 2.0 v./hr./v. A hydrogen to naphtha ratio of 1000 to 2000 s.c.f./ b. is satisfactory. The reaction is carried out with the feed in essentially the liquid phase. Temperature control can be maintained by cooling with treat gas or recycled product.
Thus the first stage is characterized by low temperature, low severity and low hydrogen consumption.
In the first stage ot hydrotreating the principal reaction is the conversion of diolens tomonooleiins to stabilize the fraction. Because of the high activity of the sullied nickel-tungsten catalyst and the low reaction ternperature, the formation of polymers and coke which would foul the catalyst is avoided. At these reaction conditions aromatic hydrocarbons are not hydrogenated. Prior art catalysts such as metallic nickel, platinum and palladium are much less satisfactory because of their sensitivity to unstable, sulfur-containing feeds. When these catalysts are used extensive feed preparation must be ernployed. The lirst stage catalyst employed herein is regenerated by conventional steam-air methods.
The partially hydrogenated fraction is passed by line 4` to the second stage hydrotreating step carried out in hydrogenation reactor 5. Pressure and temperature adjustments are made as required by conventional means. If suicient hydrogen is notvpresent in the ellluent passing through line 4, supplemental fresh or recycle hydrogenis added as necessary by means not shown. Suitable conditions for the second stage hydrogenation depend in part on the degree of saturation required and the quantity of sulfur which must be removed. Because the feed has been stabilized in the rst stage higher temperatures, more severe conditions and conventional catalysts can be used. Temperatures ranging from 400 to 700 F., pressures ranging from 200 to 1000 psig., space velocities ranging from 0.5 to 5.0 v./hr./v. and hydrogen rates of 500 to 1500 s.c.f./ bbl. are satisfactory.
Suitable second stage hydroiining catalysts comprise a hydrogenating component in which the metal is selected from the group consisting of Groups VI-B and VIII-B of the Periodic Table and mixtures of these metals distended on la suitable support material. Such metals as platinum, palladium, nickel, tungsten, cobalt and molybdenum are effective hydrogenating components and they are used in elemental form or as salts such as oxides, sulfates, nitrates, etc. The most preferred metals are nickel, tungsten, cobalt and molybdenum and mixtures thereof. Specilic examples include reduced nickel, nickel tungstate, nickel molybdate and cobalt molybdate. Suitable support materials include alumina, clays, bauxite, kieselguhr, molecular sieves, silica promoted alumina, silica promoted magnesia, charcoal, etc. The catalyst is preferably sulfided prior to use or in situ with H28 or CS2. The most preferred catalyst is sulded cobalt molybdate on silica stabilized alumina because it functions well in the saturation of olefns and in sulfur removal. The employment of this catalyst may be a critical requirement for continuous operations lwhen low sulfur aromatic hydrocarbons are being maximized as products. Aromatics are not hydrogenated to any substantial extent.
Stabilized product reduced in dioletins, olefins and sulfur is recovered by line 6. The product is passed by line 6 to separator 7. The gas fraction is removed overhead by line 8. All or a portion of the gas is recycled to the iirst stage hydrotreating reactor via lines 8, 9, 3 and 1. A portion of the gas can be purged from the system by line 10. If desired the gas can be treated to remove HES and other impurities by means not shown. The liquid product is passed by line 11 to stripper 12. Preferably a portion of the liquid is recycled via lines 13 and 1 to the first hydrotreating stage or by lines 13 and 14 and 4 to the second hydrotreating stage. Heater 15 is employed to raise the temperature of the recycle liquid to that desired in second hydrotreating stage.
From the stripper, light ends are removed overhead by line 16. A liquid fraction is passed by line 17 to fractionator 18 The desired product, i.e. a full boiling range aromatic motor gasoline is removed from the process by line 19. A fraction containing heavyends and polymers is removed from the process by line 20.
The ability of the process of the invention to handle high sulfur, high residue (gum), highly oletinic feedstocks is illustrated by the following example:
EXAMPLE 1 A naphtha (C5 to 435 F.) recovered from steam cracking to produce ethylene was used as feed. Its properties are shown in col. l of Table I.
This poor quality feedstock with no pretreatment of any kind was fed to a charge of catalyst which had been in service for 24 days. Enough steam cracked naphtha of this and other types had been passed over this nickel tungsten on alumina catalyst to evolve hydrogen sulde equivalent to over 20% of the weight of the catalyst. A metallic nickel catalyst would have beencompletely deactivated and" the precious metal catalyst seriously impaired by this amount of sulfur.-In addition, the .heavy ends, i.e. 435 F. boiling point and the 327 mg./ 100 ml. of existant gum` also would have seriously further irnpaired the activity of the precious metal catalyst. Nevertheless,.this sullided nickel tungsten catalyst operating at 279 F. inlet temperature, 623 F. outlet, 800 p.s.i.g. pressure, 0.5 v./hr./v. of fresh feed, 0.17 v./hr./v. of product recycle and 1500 s.c.f./b. of 70% hydrogen gave the results shown in the last two columns of Table I.
TABLE I Feed Total Redistitilled Product Total *Approximately equal to wt. percent conjugated diolens.
The bromine number of the feedstock was reduced from 78.9 to 31.2. The gum content was reduced from 327 mg./ 100 ml. It is evident that the high sulfur and high existant gum content of the highly oleiinic feed did not prevent the nickel tungsten catalyst operating at these conditions from eliminating to 95% of the gum forming diolens. The total product is entirely satisfactory for charging to the customary second stage operation for re'- ducing bromine number to the level of 5 or less for recovering aromatic solvents by extraction, or to whatever level is requiredr for blending in the motor gasoline pool.
In this example, the catalyst of the second stage is cobalt molybdate. When the product of the iirst stage of hydroiining is treated in the second stage at the temperatures and pressures recited previously and bromine number will be reduced to less than 5 and the sulfur content will be reduced to less than l p.p.rn. This material makes an excellent feedstock for the recovery of aromatic hydrocarbon fractions and speciiic aromatic hydrocarbons such as benzene. The second stage catalyst is regenerated as required by conventional methods.
FIGURE 2 discloses an embodiment in which the process is operated to produce both motor gasoline and an aromatic hydrocarbon fraction comprising benzene, toluene and xylene. A raw steam cracked naphtha fraction like the naphtha shown in Table I boiling in the range of C5 to 430 F. is passed by line 1 to the irst stage hydroiiner designated by reference numeral 2. A hydrogen containing gas comprising fresh treat gas and recycle gas is passed by lines 3 and 1 to the hydroliner. Hydrotining is carried out in the same manner and With the same conditions mentioned with respect to FIGURE l. Patrially hydrofined product is recovered by line 4. From the separator a gas phase is removed overhead by line 6. A portion of the gas is vented from the system by line 7. The remaining gas is recycled with or without purification by known means through lines 8, 3, and 1. Fresh treat gas is supplied by line 9. The partially treated fraction is then passed by line to stabilizer 11. A portion of this liquid is recycled via line 12. A light ends fraction is recovered from the stabilizer by line 13. Stabilized partially hydrogenated liquid is passed by line 14 to fractionator 15. In fractionator 15, a stabilized fraction boiling in the range of about 100 F. IBP to about 160 F. is removed overhead by line 16 and passed to the motor gasoline pool. The end point of this fraction is determined by whether the process is being operated to maximize benzene or to maximize the entire BTX fraction. The stabilized fraction boiling in the range of from about 300 F. IBP to about 430 F. EP is removed by line 17 as a side stream from the fractionator and this material is blended with the motor fuel components in line 16. The fraction comprising heavy ends and polymers boiling above about 430 F. is removed from the process by line 18.
The benzene G60-200 F.) or BTX (16C-300 F.) fraction is passed by line 19 to the second stage hydrofining step designated by reference numeral 20. Second stage hydrofining is carried out with catalysts and conditions similar to those disclosed with respect to the description of FIGURE 1. Conditions are adjusted to reduce the sulfur concentration in the aromatic fraction to less than about 2O ppm. preferably less than 1-5 ppm. The hydroned and hydrodesulfurized atomic fraction is passed by line 21 to separator 22. A gas fraction is recovered overhead from the separator by line 8. This gas is recycled to the first stage hydrofiner with or without purification. The hydrogen free fraction is passed by line 23 to stripper 24. A portion of this fraction can be recycled by line 26. A light ends fraction is recovered for use as fuel by line 25. The concentrated aromatic liquid stream is passed by line 27 to solvent extraction zone 28. Extraction is carried out at conventional conditions with solvents selected for aromatic hydrocarbons such as alkylene glycol, dimethylsulfoxide and sulfolane. Udex extraction with a mixture of alkylene glycol and water is the preferred extraction technique. A raffinate stream comprising saturated hydrocarbons is recovered overhead by line 29 and this material is blended with the other motor gasoline components in line 16. If desired all or a portion of the raffinate stream can be recovered by line 30 for recycle to the steam cracker. After conventional solvent recovery by means not shown the aromatic stream is passed to fractionator 32 for recovery of individual aromatic hydrocarbons. Benzene is recovered by line 33. Toluene is recovered by line 34 and xylene is recovered by line 35. The recovery scheme shown is generalized and in practice a series of towers are used. A fraction boiling above the boiling point of xylene is recovered as bottoms from the fractionation zone by line 36 and this material is mixed with the motor fuel components in line 16.
In another embodiment of the invention the liquid aromatic hydrocarbon fraction recovered by line 27 is passed to a dealkylation reactor. Benzene is maximized by conversion of higher boiling alkyl aromatics to benzene. Dealkylation may be in a thermal process or in a catalytic process. The preferred process is disclosed by U.S. Patent 3,256,357, issued June 14, 1966. The patent discloses dealkylation of a feed comprising alkyl substituted aromatic hydrocarbons such as xylene and toluene in the presence of one to six moles of hydrogen at a temperature in the range of 1050 to 1350 F. Dealkylation can also be car ried out in the presence of a catalyst. Suitable dealkylation catalysts include chromium oxide, vanadium oxide, molybdenum oxide, cobalt oxide, nickel oxide, and tungsten oxide or mixtures thereof on a suitable support such as alumina or alumina containing silica and potassium oxide. A number of hydrodealkylation processes of both the non-catalytic and the catalytic type are disclosed in the text Advances in Petroleum Chemistry and Refining, vol. 9, 1964, Interscience Publishers at pages 46-95.
The processes described herein provide an efficient, low cost means for upgrading the increasingly large quantities of raw unstable liquid coproducts derived in steam cracking for ethylene. The first stage of the process is unique in that the feed is stabilized at low reaction temperatures without excessive hydrogen consumption. Expensive pretreatment or after-treatment to remove sulfur is eliminated. When the aromatics are not separated, the stable, sulfur-free gasoline from the process has an octane number of over 100. Benzene from the extraction or dealkylation steps is of a quality suitable for cyclohexane production.
1. A process for upgrading raw steam cracked naphtha comprising the steps of (a) contacting said naphtha in substantially the liquid phase with a hydrogen containing gas in a first hydrogenation stage at mild conditions including an average reaction temperature of from about 225 to about 500 F. in the presence of a catalyst comprising sulfided nickel tungsten on a support material to produce a partially hydrogenated naphtha substantially free of diolefins,
(b) passing said partially hydrogenated naphtha to a second hydrogenation stage,
(c) contacting said partially hydrogenated naphtha in a second hydrogenation stage at conditions more severe than those of said first stage with a hydrogen containing gas and a catalyst comprising at least one metallic component to selectively convert the majority of the alkenes to alkanes, and
(d) recovering a stabilized naphtha fraction.
2. Process according to claim 1 in which said raw naphtha feed has an olefin content of 10-20 wt. percent, a diolelin content of 1-20 wt. percent and a bromine number of at least 50.
3. Process according to claim 1 in which the inlet temperature to the first stage is less than about 290 F.
4. Process according to claim 1 in which the AT across the catalyst bed in the first reactor is in the range of 60 to 400 F.
5. Process according to claim 1 in which the metallic component of the second stage catalyst is selected from the group consisting of nickel, tungsten, cobalt, molybdenum and mixtures thereof.
6. Process according to claim 1 in which the second stage catalyst is sulfided cobalt-molybdate in a support material.
7. A process for the recovery of an aromatic hydrocarbon fraction from a raw steam cracked naphtha fraction containing aromatic hydrocarbons, alkenes, alkadienes and sulfur comprising the steps of (a) contacting said naphtha in substantially the liquid phase with a hydrogen containing gas in a first hydrogenation stage at mild conditions including an average reaction temperature of from about 225 to about 500 F. in the presence of a catalyst comprising sulfided nickel tungsten on a support material to produce a partially hydrogenated naphtha fraction substantially free of diolens and gum,
(b) fractionating the effluent of said first stage to recover a fraction boiling in the range of from about 10G-430 F. containing olefins, aromatic hydrocarbon and sulfur,
(c) contacting the fraction recovered from step (b) in a second hydrogenation stage at a temperature above 400 F. with hydrogen containing gas and a catalyst comprising at least one metallic component selected from the group consisting of cobalt, molybdenum, nickel and tungsten to selectively convert the alkanes to alkenes and to remove sulfur without saturating the aromatic hydrocarbons and (d) recovering an essentially sulfur free aromatic hydrocarbon fraction.
8. Process according to claim 7 in which the fraction recovered in step (d) is extracted with a solvent selective for aromatic hydrocarbons.
9. Process according to claim 7 in which the fraction recovered in step (d) is passed to a dealkylation zone and dealkylated in the presence of hydrogen producing high purity benzene.
10. Pocess according to claim 7 in which the second stage catalyst is sulded cobalt molybdate on a support comprising alumina and 1-10 wt. percent silica.
References Cited UNITED STATES PATENTS 2,953,519 9/ 1960 Bereik et al 208-143 5 3,116,233 12/1963 Douwes et al 208-143 3,215,618 11/1965 Watkins 208-143 3,310,485 3/1967 Bereik et al 208-143 3,388,056 6/1968 Lewis 208-143 3,394,199 7/1968 Eng et al 208-143 HERBERT LEVINE, Primary Examiner U. S. Cl. XR. 208-67, 143
|Cited Patent||Filing date||Publication date||Applicant||Title|
|US2953519 *||Dec 16, 1957||Sep 20, 1960||Gulf Research Development Co||Start up procedure for catalytic hydrogen treatment of hydrocarbons|
|US3116233 *||Dec 30, 1960||Dec 31, 1963||Shell Oil Co||Low-temperature selective hydrogenation of dienes|
|US3215618 *||Sep 9, 1963||Nov 2, 1965||Universal Oil Prod Co||Hydrorefining of coke-forming hydrocarbon distillates|
|US3310485 *||May 4, 1964||Mar 21, 1967||Gulf Research Development Co||Hydrogenation of olefinic gasoline|
|US3388056 *||Aug 19, 1966||Jun 11, 1968||Exxon Research Engineering Co||Process for the hydrogenation of steam cracked naphtha|
|US3394199 *||Feb 20, 1961||Jul 23, 1968||Exxon Research Engineering Co||Hydrocarbon conversion process|
|Citing Patent||Filing date||Publication date||Applicant||Title|
|US3919074 *||Aug 22, 1974||Nov 11, 1975||Universal Oil Prod Co||Process for the conversion of hydrocarbonaceous black oil|
|US3951780 *||Oct 25, 1974||Apr 20, 1976||Exxon Research And Engineering Company||Aromatic oils by thermal polymerization of refinery streams|
|US4167533 *||Apr 7, 1978||Sep 11, 1979||Uop Inc.||Co-production of ethylene and benzene|
|US4215231 *||May 29, 1979||Jul 29, 1980||Uop Inc.||Co-production of ethylene and benzene|
|US4235701 *||Mar 30, 1979||Nov 25, 1980||Atlantic Richfield Company||Aromatics from dripolene|
|US5904838 *||Apr 17, 1998||May 18, 1999||Uop Llc||Process for the simultaneous conversion of waste lubricating oil and pyrolysis oil derived from organic waste to produce a synthetic crude oil|
|US20080194900 *||Dec 9, 2005||Aug 14, 2008||Bhirud Vasant L||Steam Cracking with Naphtha Dearomatization|
|US20150071836 *||Aug 22, 2014||Mar 12, 2015||Phillips 66 Company||Systems for pyrolysis vapor upgrading|
|EP0024299A1 *||Jul 15, 1980||Mar 4, 1981||BASF Aktiengesellschaft||Process for obtaining aromatic compounds from pyrolysed petrol|
|WO2006063201A1 *||Dec 9, 2005||Jun 15, 2006||Bhirud Vasant L||Steam cracking with naphtha dearomatization|
|WO2015061038A1 *||Oct 8, 2014||Apr 30, 2015||Uop Llc||Pyrolysis gasoline treatment process|
|U.S. Classification||208/57, 585/319, 208/48.0AA, 585/841, 208/48.00R, 208/67, 208/48.00Q, 585/265, 208/143|
|Cooperative Classification||C10G65/06, C10G2400/02|