Dehydrogenation of Ethylnaphthalene to Vinylnaphthalene
Engineering Process
development I
J. E. NICKELS AND B. B. CORSON M E L L O N INSTITUTE, PITTSBURGH 13, P A .
Vinylnaphthalene is a potentially valuable homolog of vinylbenzene (styrene). The object of the investigation was to develop a catalytic dehydrogenation process for producing it from ethylnaphthalene similar to that used in the manufacture of styrene from ethylbenzene. Under proper conditions the yield of vinylnaphthalene from ethylnaphthalene is essentially the same as that of styrene from ethylbenzene and of a-methylstyrene from isopropylbenzene-i.e., an ultimate yield of about 90%. Nineteen catalysts were tested at 600" to 675" C., 0.2 to 2 seconds' contact time, and atmospheric and subatmospheric pressure, in both the presence and absence of diluent steam. In general, the yield of vinylnaphthalene was benefited by low temperature, short contact time, and high ratio of diluent to ethylnaphthalenein the feed stock. The commercial equipment used in the manufacture of styrene from ethylbenzene is suitable for the catalytic dehydrogenation of ethylnaphthalene to vinylnaphthalene.
INYLNAPHTHALENE, a potentially valuable homolog of styrene, has usually been prepared by polystep synthetic routes (1, 6, 0-11, IS) rather than by direct dehydrogenation of ethylnaphthalene. Brief descriptions of the catalytic dehydrogenation of ethylnaphthalene have been reported-dehydrogenation over mixed copper-iron-nickel oxides on pumice and cerium oxide on pumice a t 10 to 20 mm. ( I d ) ; over 50% chromia-t%yo magnesia and 10% zinc oxide-90% alumina at 12 t o 33 mm. with nitrogen &s diluent (16); and over chromia-alumina-copper at 10 to 40 mm. (4). The major object of this work was to study the dehydrogenation of ethylnaphthalene over a considerable number of catalysts under operating conditions comparable to those used commercially in the dehydrogenation of ethylbenzene to styrene. Thus, the majority of the experiments were made a t atmospheric pressure in the presence of diluent steam. Some runs were made a t 100 mm. pressure in the absence of diluent, and a few runs were made a t 400 mm. in the presence of diluent. Under appropriate conditions the yield of vinylnaphthalene from ethylnaphthalene is essentially the same as that of styrene from ethylbenzene. I n general, the yield of vinylnaphthalene is benefited by low temperature, low pressure, short contact time, and high molal ratio of diluent to ethylnaphthalene. APPARATUS, PROCEDURE, AND FEED STOCK
The apparatus for operation a t atmospheric pressure was essentially the same aa that previously described (8,14). The reactors were usually of Vycor (19-mm. inside diameter), the vaporizerpreheater being an integral part of the reactor and packed with 4-8 mesh Vycor chips. I n one run the reactor wa8 a 2-inch copper-lined steel tube (4.5 feet long) equipped with an integral
vaporizer-preheater section packed with brass jack chain. The apparatus for operation under reduced pressure is shown diagrammatically in Figure 1. A Cenco Pressovac pump was found to be well suited for reduced pressure operation in that it had sufficient capacity and permitted gas collection on the pressure side of the pump. The pressure within the dehydrogenation system was maintained a t the desired level by means of a manostatcontrolled solenoid valve. Thus, it was possible to collect airfree gas samples which cannot be done in an air-leak controlled reduced pressure system. The operational procedure was essentially the same as that followed in the dehydrogenation of ethylbenzene (8) and isopropylbenzene (14). The constants of the ethylnaphthalene feed stock (approximately 50-50 mixture of alpha and beta isomers) were: boiling point, 257" to 259" C.; ngo, 1.6022; dzo, 0.9997. Fresh feed was used in all the quantitative runs. ANALYSIS OF PRODUCTS AND CALCULATION OF YIELDS
The duration of the runs ranged from 3 to 144 hours. I n the case of the 3-hour runs the catalymte from each hour's production was analyzed; in the 30-hour runs the composition of the product from each 4-hour period was determined; with the longer runs the product from each 12- or 24-hour period was analyzed. Table I presents averaged values for the entire runs. The yield per pass of vinylnaphthalene was defined as grams of vinylnaphthalene produced per 100 grams of ethylnaphthalene processed; the ultimate yield of vinylnaphthalene, as grams of vinylnaphthalene produced per 100 grams of ethylnaphthalene consumed. The sum of the per pass and ultimate yields (socalled yield index) furnished a rough relative measure of the dehydrogenating activity of a catalyst under the t a t conditions employed. The vinylnaphthalene content of the liquid catalyzate !\-as calculated from the bromine number. The naphthalene and methylnaphthalene contents of the catalyzate were calculated from the amount and composition of the gas, assuming ethylene and ethane to correspond to naphthalene, and methane to methylnaphthalene. This was the method employed by Mavity, Zetterholm, and Hervert ( 7 ) in estimating the benzene and toluene contents of catalyzate from ethylbenzene dehydrogenation. The authors verified the applicability of the method in several instances by distilling liquid catalyzate (inhibited by 0.57, of p-tert-butylcatechol) at 10 mm. through a 25-plate glass column packed with glass helices and identifying the plateau fractions. The values for naphthalene and methylnaphthalene obtained by the two methods (gas - analysis and distillation) did not differ bv more than 0.8%. The amount of high-boiling residue in the liquid catalyzate was determined by topping a 25-gram sample under an absolute pressure of about 0.7 mm. with the distilling tube at 80° C. and the condensing tube at -78" C. The ethylnaphthalene content of the catdyeate (and from that, the ethylnaphthalene consumed) was determined by difference.
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Figure 1. Apparatus for Dehydrogenation under Reduced Pressure A. E. C. D. E. F.
G. C e n c o Pressovac pump Fe& s t o c k reservoir H. G a s s a m p l i n g valve Bellows p u m p Dehydrogenation f u r n a c e I. Off-gas bubbler Water-cooled c o n d e n s e r . I .W e t test meter Ice-cooled t r a p K. Differential manometer L. Open-end manometer M a n o s t a t - o p e r a t e d solenoid valve M. A d j u s t a b l e manostat
The gas was analyzed by conventional absorption-combustion technique and the carbon on the used catalyst was determined by combustion. QUALITATIVE ISOLATION OF @-VINYLNAPHTHALENE. Separe tion of the vinylnaphthalene isomers from liquid catalyzate by distillation was not successful because of excessive polymerization even when operating a t a pot temperature (approximately 110'C.) corresponding to distillation a t 5 mm. pressure; polymerization inhibitors were ineffective a t this temperature. Continuous distillation as contrasted t o batch distillation would have decreased the amount of polymerization. However, crystallization afforded a means of partially separating the beta isomer from liquid catalyzate when its concentration was sufficientlyhigh. This was accomplished by flash-distilling liquid catalyzate at 0.5 to 1 mm. pressure in a continuous unit to remove dark-colored, highboiling components; cooling the distillate to about -45' C. with stirring to produce a thick slurry of @-vinylnaphthalene crystals, and processing this slurry through a precooled centrifuge to give a filter cake of crude product which was recrystallized from methanol or petroleum ether to give a product melting a t 66" to 66.5' C. (dibromide: melting point, 85.5' t o 86' C.). Analysis of B-Vinylnaphthalene Calcd. for C~zHio: C 93.46. H 6 . 5 4 , Br No. 1 0 3 . 6 Found: C: 93.29; H: 6 . 6 1 ; Br No.: 1 0 2 . 3
PRODUCTION OF @-VINYLSAPHTHALENE B Y RECYCLLNG. In a few qualitative experiments, mother liquor from the above crystallization was recycled through the dehydrogenation and crystallization steps until the greater part of the original @-ethylnaphthalene content of the feed had been converted to, and removed as P-vinylnaphthalene, leaving a final mother liquor considerably enriched in alpha isomers. This alpha-rich mother liquor was hydrogenated to an alpha-rich ethylnaphthalene mixture which was then catalytically isomerized to a beta-rich mixture suitable for further dehydrogenation and recovery of pvinylnaphthalene. CATALYSTS
Nineteen catalysts were tested; some were granular (approximately 4-8 mesh) and others were in pellet form ( l / 8 X l / g inch). Several catalyst preparative methods were employed-coprecipitation, impregnation, physical admixture, sulfidation, and calcination (of natural carbonates). Table I1 lists the catalysts in the order of decreasing dehydrogenating activity as indicated by yield index; pertinent preparative data are included.
Inasmuch as this study was of the nature of an exploratory survey, no attempt was made to establish optimal conditions for any of the nineteen catalysts tested. The general trends, as functions of operation and catalyst, are briefly discussed below (see also Table I). A plot of ultimate yield versus EFFECT OF TEMPERATURD. temperature reveals the detrimental effect of increase in temperature. For example, a t 750 mm. and a diluent ratio of 10 to 1, the vinylnaphthalene yields were decreased from 69.3 and 75.6% (runs 2-D, 5-E) a t 600' C. to 50.6 and 71.9% (runs 2-F, 5-B) a t 650' C., respectively. At 400 mm. and a diluent ratio of 5 to 1 the yield was decreased from 78.7% a t 600" C. (run 2-C) to 56.5% a t 650' C. (run 2-E). Although no runs were made in which EFFBCT OF PRESSURE. pressure alone was varied, data from runs in which both pressure and diluent ratio were varied show that increase in pressure lowered the ultimate yield of vinylnaphthalene. A plot of yield versus diluent ratio-pressure combinations of 0 to 1 a t 100 mm., 5 to 1 a t 400 mm., and 10 to 1 a t 750 mm. reveals that the yield dropped as the pressure rose despite the simultaneous increase in diluent ratio (demonstrated to be beneficial). For example, the yields obtained in runs 2-B, 2-C, and 2-D a t 600" C. with diluent ratio-pressure combinations of 0 t o 1 a t 100 mm., 5 t o 1 a t 400 mm., and 10 to 1 at 750 mm. were 84.7, 78.6, and 69.3%, respectively. Data from companion runs 2-E and 2-F, 6-B and 6 C , &A and 8-B illustrate the same relationship.
TABLE11. CATALYSTS Catalyst No. 1 2 3 4
5
6
7 8 9
10 11
12 13
14 15 16
17 18 19
Composition Calcined magnesite
Preparation or Source Crushed magnesite calcined at 800° C. for 6 hours, pilled 72.4% MgO-18.4% FezOa- Reference (6) 4.6% Kz0-4.6% CUO 15% C r * 0 ~ - 8 5 7AlzOa ~ Activated Ah03 pills impregnated with aq. CrOa, dried a t 110' C., calcined a t 600° C. ( 8 ) 1.3y0 Cu0-98.7% AlzOs Activated A1208 pills impregnated with aq. Cu(NOa)z, dried a t 110' C., decomposed a t 600' C. Calcined dolomite Crushed dolomite (0.349 SiOz, 0.05 FezOs, 0.09% AliOa 80.83% C a z 21.17% MgO, 47.5261, ignition loss) calcined at 650' C. for 4 hours 60% Zn0-30.4% AlzOa- Aq. paste of ZnO, Al(OH)s, Ca(OH)%, 9.6% CaO dried, pilled l.2YONi0-98.8y0 AlzOs Activated AlzOa pills impregnated with aq. Ni(NOs)z, dried a t 110' C., decomposed at 650° C. ( S ) ZnCOs Precipitated from aq. ZnSOr b y NanCOa filtered washed, dried a t 110d c., pilled 60% CuCOs-50% ZnCOa Precipitated from aq. ZnSO4-CuSOa b y NazCOz filtered, washed, dried at 110" C.,'pilled 1.5% NiS-98.5% AlsOa Ni0-AIzOa catalyst pilla sulfided at 650' C. by HzS 92.5y0 calcined dolomite- Catalyst y o . 5 impregnated with aq. 7.5% CrzOs CrOs, dried a t 110' C., calcined a t 600' C. CeeOa Cerium oxalate decomposed over Meeker burner t o brown powder, pilled 25% CuCOa-75% ZnCOa Precipitated frsm aq. ZnSOa-CuSOa b y NazCOa filtered, washed, dried a t 110' C.,'pilled A1203 Activated AlzOa pills 7.7% NiS-92.3% AlaOa Ni0-AlzOa catalyst pills sulfided a t 650' C. b y HnS 9 % MoOa-91% AlzOs M . W. Kellogg Co. Calcined limestone Reference (9) Bauxite Porocel. 3 5 2 7 Ti02 1 7 0 % FezOa, 74.55% AlzOaP 15.43% 'SiOz, 4.80% ignition loss 54% ZnO-27.3% AlzOa- Aq. paste of ZnO,. Al(OH)a, KzCOs, lOvoKzCOa-8.7% CaO Ca(OH)z, dried, pilled
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contributory variables. Table I11 lists the catalysts in order of decreasing ac-
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tivity for several sets of operating conditions. XINOR EFFECTS OF CATALYST Cox3 Y POSITIOK. Calcined dolomite (run 5-R) 60 \vas better than calcined dolomite imIn! pregnated with 7.5% chromia (run $50 D 11-A)-i1.9~, yield versus 66To yield, , w 2 7r e s p e c t i v e l y . C a t a l y s t 10 (l.57c n i c k e l sulfide-98.57, alumina) was / o o ! g L better than catalyst 15 (7.7% nickel 30 sulfide-92.37, alumina)-707, yield (run t. I 10-A) versus 67.27, yield (run 15-11). > i Catalyst 9 (507? cupric carbonate-507, HOURS zinc carbonate) was slightly better than Figure 2. Data from 120- and 144-Hour R u n s catalyst 13 (25% cupric carbonate75% zinc carbonate)-62.2% yikld (run 9-A) versus 60.77, yield (run 13-A). The arrangement of the three impregT A B L E 111. ULTII\lBTE YIELD O r ~ I S Y L N A P H T H A L E N EUS. CAT.4LYST AND 0 P E R h T I X . G COXDITIOXS nated alumina catalysts in order of decreasing activity is catalysts 4, 10, and (GOO" C., 100 h l m . . S o Diluent) (650O C., 750 Rlin., HzO: E S a = 10: 1) CataCata7; the yields being 72.3, 70 and lyst Yield, lyst Yield, No. Catalyst % KO. Catalyst % 68.67, (runs 4-B, 10-A, and 7-B), re1 Calcined magnesite 90.3 4 CuO-AlzOa 72.3 spectively; the impregnants (equimolal) Ni0-rllzOs 86.6 5 Calcined dolomite 71.9 being 1.370 of cupric oxide, 1.5% of Mg0-FezOa-Kz0-Cu0 84.7 1 Calcined magnesite 71.6 6 ZnO-Al20s-CaO 83.9 3 CrzOa-AlzOa 70.0 nickel sulfide, and 1.27, of nickel oxide, 4 CuO-AizOs 82.2 SiS-Al203 70.0 XO-AlzOr 68.6 respectively. The addition of potassium 3 Cr20s-dIz03 81.6 lo 7 79.7 12 NiS-Al203 14 A1208 67.2 carbonate to the combination zinc z6.6 14 $1203 66.7 12 CeaOa 16 IbloOa-AhOa 14.1 17 CaCOs 66.5 oxide-alumina-calcium oxide was not 18 Bauxite 70.7 11 Calcined dolomite-Crz0a 66.0 64.6 beneficial-catalyst 19 (64% zinc oside8 ZnCOs 6 (650' C., 400 R l m . , HaOCEXO = 5 : 1) Zn0-AIzOa-CaO 62.2 27.37, alumina-107, potassium car19 Zn0-AlaO8-~hCOa-Ca0 60.2 Catalyse YieId, 2 MgO-Fe203-KzO-CuO 20.6 bonate-8.77, calcium oxide, run 19-A, NO. Catalyst % yield 60.296) versus catalyst 6 (60% zinc 6 8 ZnO-Al20a-CuO ZnCOs 70.7 79.5 oside-30.47, alumina-9.6Yo calcium 2 MgO-Fe~O~-XzO-CuO 66.5 oxide, run 6 4 , yield 62.27,). a E S = ethylnaphthalene. LONGERRUNS. Catalyst 5 (calcined dolomite, run 5-C) was operated for 120 hours (19-mm. inside diameter Vycor The ultimate yields of vinylnaphthalene a t 600' C . and 100 tube, 750 mm., 650' C., steam-ethylnaphthalene molal ratio mm. in the absence of steam ranged from 70.7 to 90.375, whereas 10 to 1) and catalyst 2 (magnesia-ferric oxide-potassium oxidea t 6.50' C. and 750 mm. in the presence of 10 moles of steam per cupric oxide, run 2-A) was operated for 144 hours (%inch inside mole of ethylnapht,haIene the yields n-ere approximately I5 to diameter copper-lined steel reactor, 750 mm., 600' C., steam207, lower, ranging from 50.6 to 72.376 (Table 111). ethylnaphthalene molal ratio 20 to 1). The yield data are EFFECTOF DILUENT.The beneficial effect of diluent is sh0TT-n presented in Figure 2. by the following exaniplcs. CONCLUSION Operation ai; 600" C., 750 mm., and I-second contact time without diluent (run 6-F) gave a 59.6% ultimate yield, whereas operThe dehydrogenation of ethylnaphthalene over nineteen ation at the less efficient temperat,ure Of 650" and the longer catalysts m.as studied; the t,emperatures ranged from 600' to contact time of 2 seconds, but in the presence of 10 moles of dil675" C . , and the cont'act times from 0.17 to 2 seconds. Reduced uent per of ethylnaphthalene, gave a 75.6% pressure of the order of 100 mm. was found t o be advantageous; (run 5-E). Operation a t 650" C . and 750 mm. with 10 moles of several ca,talysts gave ultimate yeilds of 81.6 t o 90.3%. Operadiluent steam gave a 68.6% ultimate yield (run 7-B), while under the same conditiolls except for the absence Of Steam a Yield of only tion at atmospheric prRssure mith steam as diluent gave ultimate 49.27, was obtained (run 7 - C ) . yields of 70 t o 81.57,, depending on the catalysts and the condiThe advantage of benzene as diluent is that i t is miscible with tions of operation. In general, the yields of vinylnaphthalene obhinable from ethylnaphthalene are of the same order as Chose ethylnaphthalene which simplified the feeding of the cllarge, but its disadvantage is that it does not separate from the hydrocarbon of styrene from et,hylhenzcne and of a-methylstyrene from iaocatalyzate as does water but must be distilled therefrom. The propylbenaene (8). data indicate steam to be preferable to benzene vapor as diluent at ACKNOW LEDGRIENT 650" C. and 750 mm. (run 5-B, 10 moles of steam, il.9Y0 yield; Thanks are expressed to Nicholas Greco, Grace Holmes, Nancy run 5 4 , 20 moles of steam. 74.40, J-ield; run 5-D, 10 moles of Konstanzer, and Ruth Leffler for assistance in the experimental benzene, 53.77, yield). xvork. The work was performed under the Koppers CO., T~lc., Operating E~~~~~OF coNTACT T ~ ~ ~ ~ : . at 6000 c., 750 mm,, and a benzene-ethylnaphthalene molal ratio of 5 to 1, the ultimultiple fellowship on tar SYntheticS. mate yield of vinylnaphthalene was increased by about 16% LITERATURE CITED (runs 3-B and 3-C) by decreasing the contact time from 0.33 t o (1) Brandis, Ber., 22, 2148 (1889). 0.17 second. (2) Corson and Webb, U. P. Patent 2,444,035 (1948). EFFECTOF CATALYST.From the data available, i t is impos(3) Ibid.. 2,470,092 (1949). (4) Ghosh and Guye, Petroleum ( L o l z d ~ f l . )13,283 , 11950). sible to establish which is the best, catalyst because of numerous 0
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(11) Sontag, Compt. rend., 197, 1130 (1933); Ann. chim., [ l l j 1, 399 (1934). (12) Suida, Austrian Patent 132,042 (1932); U. S. Patent 1,985,844 (1 934). (13) TVhmu~ and Daudel, Compt. rend., 147,678 (1908). (14) Webb and Corson, IND. ENG.CWEM., 39, 1153 (1947). (16) Zal‘kind and Arbueova. Plasticheskie Massy, Sbornilc State%, 3 , 2 4 9 (1939).
(5) Johnson, U. S. Patent 2,468,759 (1949). (6) Kearby, I b i d . , 2,395,875 (1946). (7) Mavity, Zetterholm, and Hervert, Trans. Am. Inst. Chem. Engrs., 41, 519 (1945). (8) Nickels, Webb, Heintselman, and Corson, IND. ENG.CHEM., 4 1 , 5 6 3 (1949). (9) Palfray, Sabetay, and Sontag, Compt. rend., 194, 2065 (1932). (10) Shoruigin and Shoruigina, J. Gen. Chem. (U.S.S.R.), 5, 555 (1935).
RECEIVED October 7, 1950.
Fractional Extraction Process for Recovery of Pure Tar Acids
Enginnering Process development
M. B. NEUWORTH, VERA HOFMANN,
AND T. E. KELLY PIlTmURGH CONSOLIDATION COAL CO., LIBRARY, PA.
Caustic soda extraction of tar acids from tar distillates has the disadvantages of batch operation and consumption of reagents. A review of the literature showed that continuous solvent extraction methods had been explored using oxygenated solvents. These processes were not too satisfactory, owing to the production of impure tar acids or the necessity of recycling a substantial fraction of the tar acids to obtain a satisfactory recovery. A process has been developed for the recovery of tar acids from low temperature tar distillates based on fractional solvent extraction with aqueous methanol and hexane. Tar acids were recovered by this process in a yield and purity comparable with caustic soda-extracted acids. This extraction process has the advantage of requiring no recycle of feed. TAW cost, easily recovered solvents are used. For tar distillates containing high concentrations of tar acids such as occur in low temperature tar, this solvent extraction process appears to be a significantly Gwer cost nperation than caustic soda extraction.
C
AUSTIC soda extraction of tar acids from tar distillates has been practiced commercially for a long time. Because the recovery of tar acids by this process is accompanied by the consumption of reagents in the extraction, springing, and regeneration cycles, reagent costs are an important item. A further disadvantage of the caustic soda process is the use of batch operations in most commercial installations. The development of a solvent extraction process for the recovery of tar acids from tar distillates was of interest in view of the economies that might be effected in making the process continuous and elimination of any reagent consumption. The application of a solvent extraction process t o low temperature tar distillates is of particular interest in view of the high concentration of tar acids. This is due to the fact that a solvent extraction process increases in relative economic atfiractiveness over the caustic soda process with increased concentration of tar acids in the tar distiIlate A number of oxygenated solvents have been described in the literature for tar acid extraction including water, methanol, ethanol, formic acid, and acetic acid (3, 6, 8, 11). The German “Metasolvan” process (9) used aqueous methanol or ethanol for the removal of tar acids from low temperature tar distillates. The object of this process was to prepare phenol-free neutral oil
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for use as a Diesel fuel. The tar acids recovered by this process contained as much as 20% neutral oil. Petroleum ether was suggested for purification of the methanol extract in a second extraction column. A recent invest,igation by Prutton et al. (10)showed that aqueous methanol was a promising solvent for tar acid extraction as a result of equilibrium distribution studies on pure phenols and hydrocarbons of the type present in coal tar. However, the data indicated that production of hydrocarbon-free tar acids is not possible by the use of a single solvent of the alcoholwater type. A similar purification of the methanol extract with a paraffinic hydrocarbon has been proposed. The successful development of a two-column, two-solvent process for the recovery of pure tar acids has the serious drawback of requiring recycle of a substantial part of the tar acids from the purification column to the extraction column as a result of carry-over into the hydrocarbon purification solvent. Accordingly, the development of a fractional extraction process for the quantitative recovery of pure tar acids in one column &as undertaken. Fractional liquid extraction offers a convenient technique for complete recovery of one component of a mixture provided an operable pair of solvents can be found. Where commercial application of the process was intended, the choice of solvents was limited by considerations such as solvent cost and ease of recovery. The authors’ investigations established that aqueous methanol of the proper concentration and a paraffinic naphtha possessed the desired properties. A paraffinic naphtha was preferred over an aromatic naphtha since i t would have a higher selectivity resulting in less carry-over of tar acids, For the laboratory studies, a multistage extraction column of the type described by Scheibel(18) was used. I n order to simplify the interpretation of results, prior to extraction of tar distillates, studies were carried out on a synthetic tar acid oil composed of 8-methylnaphthalene and m-cresol. This mixture represents the two major classes of compounds in tar distillates-namely, tar acids and neutral oil. This particular mixture was selected on the basis of the presence of the two components in the tar distillates under examination, availability of the individual compounds in high purity, convenience in handling, and ease of analysis. It was of interest to clarify certain extraction variables on this twocomponent system before attempting the extraction of the complex tar distillate mixture. The effect of methanol concentration, motor speed, choice of continuous phase, column throughput, and optimum solvent ratios were of prime concern,
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