DEVELOPMENT OF HOUDRY DETOL PROCESS A .
H. W E l S S A N D
LEE FRIEDMAN
Hoiidry Process and Chemicals Co., Philadelphia, Pa.
The recent development of hydrodealkylation for converting alkyl aromatics to unsubstituted aromatics has opened a new field of petroleum technology. The catalytic Houdry Detol process was developed utilizing nitration grade toluene as a charge stock for producing high purity benzene. The process is also applicable for dealkylation of aromatics of higher molecular weight than toluene, as well as substituted naphthalenes. Laboratory studies leading to the commercialization of the Detol process are summarized. Both isothermal and adiabatic pilot unit data are cited. The laboratory data were evaluated kinetically by programming the reaction model on a digital computer. As a result, the experiments required to characterize the process variables were minimized, and the process was developed in a period of six months.
ITH THE AD\'EST
of petroleum deribed toluene, a n active
W commercial interest in hydrodealkylation processes to produce benzene from toluene has developed. The Houdry Detol process is a catal>tic technique for achieving this goal. The basic reaction is CbH,CH,
+ H! * CEH6 + CH,
(1)
The reaction is effectively unidirectional, and it is rate rather than equilibrium controlled. There has been considerable w'ork in the h) drodealkylation field, and an excellent bibliography of literature and patent references is provided in an article by Stijntjes and others ( 6 ) . The Detol process in its commercial application a t the Croxvn Central Petroleum Corp. refinery. Houston, Tex., is described in an article by \\'eiss et al. (70). A review article describing the mechanisms of thermal toluene demethylation has also been published (9). Thermal Reactions
Silsby and Saivyer ( 5 ) found that the homogeneous thermal toluene demethylation reaction in the presence of excess hydrogen is a 1.5 order reaction. The reaction proceeds by the following steps: Hydrogen dissociates into atoms directly:
followed by: C6Hj.CH3 CsHj,
+ He + H2
+
+
C6Hj.
CsH6
+ CHJ
+H.
(3)
(4)
Reaction 3 is analogous to the conversion of para to ortho hydrogen : (5)
Silsb) and S a n ) e r ( 5 ) measured in a flow system the value of the activation energy for homogeneous thermal hydrodea1k)lation of toluene to benzene to be 50 5 kcal. per gram-mole. The frequency factor for the reaction is 10'a.6= (mole/liter) per second. Values of the activation energy ranging from 43 to 54.5 kcal. per mole have been reported by Matsui ( 4 ) and Tsuchiya (8). i2n activation energy of 56.0 kcal. per mole and a frequency factor of 5.1 X 10" (mole/ liter)-'I2 per second are reported by Stijntjes ( 6 ) . I n addition to the major hydrodealkylation reaction whose mechanism is represented by Equations 2, 3, and 4, termination reactions occur to a slight extent in the q s t e m . Heavy residues form, and examples of these are:
*
2 CsHjCH2. * C ~ H ~ C H ~ C H ~ C B H , 2 C6H.j. + CtjHj.C6Hj
Equations 3 and 10 illustrate the formation of bibenzyl and biphenyl, respectively. Biphenylmethane can also form by an analogous route. In practice, small quantities of all of these compounds, as well as their methyl substituted homologs, are found. Catalytic Hydrodealkylation. T h e reaction in a catalytic hydrodealkylation system is conducted a t elevated pressure and temperature. Small amounts of ring hydrogenation can occur, as illustrated by:
T o produce benzene that is not contaminated v i t h cyclohexane, it is necessary to operate a t conditions of sufficient severity that the formation of cyclohexanes is negligible. If such compounds are formed, or if paraffins are present in a charge stock to a hydrodealkylation unit, they are removed by the hydrocracking reaction which occurs simultaneously a t the hy drodealkylation operating conditions. The major hydrocracking products in the Detol system are ethane and propane; because of the temperature level of the catalytic reaction little methane is formed. Some hydrocracking reactions are :
+
C ~ H I I C H ~3 H ? -P C,HB where the dissociation of a hydrogen molecule is a true equilibrium reaction. This phenomenon explains the 1.5 order kinetics for toluene h>drodealkylation, and the rate expression for Equation 1 can be uritten exactly analogous to Equation 7 . -d[C6H6'CHJ1 =
dt
d[Cdt 6H,l =
kT[CGH,.CH3][H2]1J2 ( 8 )
(9) (10)
CiHltj
+ 2 C2H6
+ 2 H2 * C3Hs + 2 C?Htj
(12)
(13)
Termination reactions and side reactions occur to only small extents in the catalytic system. For this reason, it is within experimental accuracy to consider only the major reaction the 1.5 order hydrodealkylation of toluene to benzene, in developing a reaction model. VOL. 2
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163
(I;- . f
i
EANDA
I
=; w=
I
CALCULATED CONVERSION
I I I
ACTUAL CONVERSION
I
I E-
Figure l .
Calculation of optimum A and E
Kinetic Equations. Using algebraic symbols, the basic rate equation for toluene hydrodealkylation (Equation 8) can be rewritten:
where CT
=
toluene concentration
C, = H2 concentration
e = empty reactor contact time kT = reaction rate constant for toluene hydrodealkylation
T h e integrated form of the equation is as follows:
where T = temperature A = frequency factor E = activation energy W = fractional toluene conversion n = pressure Hi,G,, B, = moles of hydrogen, gas, and benzene a t inlet per mole of toluene Application of Equations. For isothermal operation, Equation 15 can be utilized directly to predict the rate constant k T for any combination of operating variables and an experimentally determined conversion. However, in practical experimentation, it is difficult to achieve completely isothermal reaction conditions, since the hydrodealkylation reaction is highly exothermic. Unless reactor diameters are very small, e.g., 0.75 inch, heat transfer from the catalyst bed to a surrounding constant temperature bath will lag, and a definite temperature profile will result in the catalyst bed, both longitudinally and laterally. To confine experimentation to a small-diameter isothermal reactor is not generally satisfactory if it is desired to obtain both precise material balance data and quantities of products sufficient to evaluate purity and by-product production results. 164
I & E C P R O C E S S DESIGN A N D DEVELOPMENT
On the other hand, the calculation of the rate constant is confounded in a n adiabatic system, since the temperature varies with contact time. The right hand side of Equation 15 cannot be graphically integrated to determine the rate constant, since E i n the expression c - ~ is / a~ nonseparable ~ variable. I n a n adiabatic experiment, temperature data are available as a function of distance through the reactors. Distance is. barring any significant pressure drop or gas volume change. directly proportional to contact time. T o establish rate constant information from adiabatic data, the relationship of Equation 15 was programmed on an IBM 650 computer. The program was arranged so that temperature data were tabulated as a function of contact time and were part of the data input. By this technique, any deviations of a n adiabatic reactor from true adiabaticity are not significant, since actual catalyst temperatures are used in the program. Tests of nearly isothermal data in Equation 15 permitted an initial approximation of the values of A and E. This initial set of values was used in the computer program to integrate numerically the right hand side of Equation 15 and then to calculate the predicted conversion for each adiabatic pilot run made. All runs of the experimental program were evaluated in this manner, and the sum of the squares of the deviations of predicted us. actual conversions for all of the runs were calculated. A systematic variation of A and E was then made. and the entire procedure repeated. By this means, sum of squares deviations corresponding to many values of A and E were obtained. The combination of A and E that resulted in the minimum sum of squares (least squares) deviation of predicted and actual conversions was chosen as the most applicable for the system. The advantage of this procedure is that the rate data were obtained under conditions that approximated commercial design. Figure 1 shoxs how the best values of ‘4 and E for toluene demethylation were chosen. A similar problem of obtaining rate data in a nonisothermal system is discussed by Towell and Martin (7). These authors used Arrhenius plots as successive approximations to arrive at the best values of A and E.
Process Development
The development of the Detol process proceeded in three distinct stages, each corresponding to a more elaborate pilot unit. While this report is concerned with the last stage-the development of the design model-a short discussion of the experimental background is of interest. Isothermal Experimentation. The initial experimentation on Detol was conducted in a small 100-cc. capacity isothermal pilot unit. Exploratory studies showed that catalyst screening could be accomplished by short runs of 12-hour duration, conducted a t 75 seconds’ contact time, 515 p.s.i.a.: and 1125 F. using once-through hydrogen as a diluent gas. In this unit, a t these conditions, thermal toluene conversion was only 12%, and any major conversion increase effected by a catalyct could be readily detected. The unit was not equipped to recycle gas, and material balances were not adequate in these short runs to obtain precise selectivity data. However, conversion figures were reproducible, and screening of catalyst formulations was accomplished quickly. When the Detol catalyst choice was fixed by the results of the screening study, sustained runs were made in the 100-cc.
unit to demonstrate that catalyst activity did not diminish with time on-stream and to obtain an initial estimate of the process selectivity. Catalyst coking data were also obtained that were useful for later coke correlations. The experimental results of the tivo sustained runs are shown in Table I. Note that neither benzene selectivity (3) nor toluene conversion decreased over the course of the runs. Selectivities greater than 100 mole 7 0 were, of course, due to the inaccuracies of material balances. Preliminary economic studies showed that a conversion level of a t least 65Y0 was required for a n economic process, and this conversion was shown to be feasible. Simultaneously with the sustained runs, experiments were made in a larger (2300 cc.) isothermal unit. This pilot unit had charge and product systems similar to the adiabatic pilot unit described below, but the reaction section consisted of two 1150-cc. isothermal reactors immersed in molten lead baths rather than four 300-cc. adiabatic reactors. The larger isothermal pilot unit was equipped to recycle gas, and except for the isothermal feature simulated a commercial system closely. Material balances from this unit could be closed \vith an accuracy of =k3 wt. 70,and selectivity data of high precision were obtained. Runs lasting from 24 to 48 hours each were made in this unit to demonstrate the effects of operating conditions, use of recycle gas, hydrogen purity, and catalyst regeneration. However, these reactors were not isothermal. Because of the severe lateral and longitudinal temperature gradients measured, valid coking and rate data could not be obtained from this unit, although product distribution and selectivity data corresponding to a given toluene conversion were found to agree with later adiabatic pilot data. The runs made in this unit are tabulated in Table 11. Catalyst regenerations were made every other run and are indicated in the tabulation. A concluding run showed no
Table 1.
Sustained Runs in 100-Cc. Isothermal Unit
Once-through hydrogen, 51 5 p.s.i.a., 75 seconds' contact time
Temp.,
Time On-Stream,
F. 1125
Toluene Conuersion,
5%
Hr. 24
47.6 47.0 49,3 64.5 67.3 65.7
96
1150
200 12 176 231
Moles C6Hs Produced M o l e C6HsCHa converted
Benrer'e Selectivity," M o l e 70 98.4 101.3 97.8 100.8 97.4 101.8
,oo,
significant change in catalyst activity from the initial run, made a t almost identical conditions. Adiabatic Experimentation. At this point experimentation was shifted from the 2300-cc. isothermal unit to an adiabatic unit. The advantages of the adiabatic unit over the isothermal unit were as follows: There were no transverse temperature gradients in the catalyst bed. Temperature data could be interpreted with confidence. The temperature profiles closely approximated those to be expected in commercial operation. Conversions, selectivity, and by-product distribution data could be transposed directly to commercial design. The unit was of sufficient size (1200 cc. catalyst capacity) to permit obtaining precise yield data. Figure 2 is a flow diagram of the unit. Charge stock for the process was petrochemical nitration grade toluene. Mass spectrometer analysis of the charge stock showed a purity of 99.99 vol. Yo toluene. Charge stock was pumped from a calibrated tank into a gas stream composed of fresh (make-up)
RECYCLE GAS J
CHARGE PUMP Figure 2.
Four-reactor adiabatic pilot unit VOL. 2
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Table II. Cumulative Hr. on Stream
24 42 27 51 24 48 24 48 24
Runs in 2300-Cc. Isothermal Pilot Unit
Contact Time, See.
Catalyst Outlet Temp,, F.
72 1131 73 1202 Catalyst Regeneration 69 1180 112 1177 Catalyst Regeneration 120 1176 38 1176 Catalyst Regeneration 72 1130 71 1200 Catalyst Regeneration 73 1119
Toluene Conz~ersion, 6: io
46.3 60.4 59.5
71.9 69.5 34.6
54.2 85.8 44.2
hydrogen and recycled high pressure flash gas. Gases \rere metered by rotameters. The combined liquid and gaseous charge was then passed through a 1/8-inch inner diameter coil immersed in a lead bath, \rhere it was heated to the appropriate inlet temperature, and then the reactants entered the first reactor. Effluent gas from the first reactor was then cooled by passage through a second stage coil immersed in a lead bath before entering the second reactor at a desired temperature. This could be repeated for a total of four reactors with three interstage coolers. Vse of either temperature staging or by-pass valving permitted operation of the unit as a one-, two-, three-, or four-reactor sysrem. Reactor effluent was condensed at 70' F., and gaseous and liquid products \cere separated by flashing. hlinor quantities of product remaining in the gas stream were condensed by subsequent cooling to 32' F. in a coil immersed in an ice bath. Gas from the ice condenser was split into a recycle gas stream and a high pressure flash gas stream which was vented through a pressure controller and meter. Liquids condensed both at 70' and at 32' F. were combined and flashed to atmospheric pressure. The atmospheric pressure flash gas was also metered. Material balances generally closed within r t 3 wt. 70. Following a run. the catalyst was either removed from the reactors and analyzed for coke deposits or regenerated in place. For the latter purpose, compressed air from the laboratory supply system was metered into the unit. Combustion gases were analyzed by Orsat techniques. Figure 3 is a cross-sectional diagram of one of the adiabatic Detol reactors. The reactor is made of 1.25-inch Schedule 40, Type 304 pipe. Electrical circuits above and bvloiv the catalyst bed prevent longitudinal heat leakage. The void space above the catalyst is filled by a stainless steel bar, or "dead man." Catalyst temperatures are measured by thermocouples located in a thermowell a t the center of the bed. \$'all temperatures are sensed by thermoblocks welded to the reactor \vall. For adiabatic operation, heat input from electrical circuits \round around the reactor is adjusted so that each reactor wall temperature matches the corresponding catalyst center line temperature. Theoretically, there is no heat flow from or input to the catalyst bed when these two temperatures are matched. I n actual practice, even with 6 inches of insulation surrounding the reactor, operation is significantly nonadiabatic. It is suspected that the precision of temperature measurement in the 1000° to 1200' F. range is one major cause of nonadiabaticity. 166
I & E C PROCESS D E S I G N A N D D E V E L O P M E N T
Figure 3. Cross-section of 300-cc. adiabatic reactor
1
Figure 4. tion
35.400 CALIWOL
-+
Arrhenius plots for toluene hydrodealkyla-
The enthalpv change for the h\ drodealkylation reaction, Equation 1. is AH = -21, 500 B.t u. per 1b.-mole of toluene converted .in\ hydrocracking of toluene to nonselective products such as ethane and propane results in additional quantities of heat being released during rhe reaction. A reproducibility of no better than 16000 B.t u. per 1b.-mole was established in the pilot plant. For this reason. it was concluded that enthalpy changes determined experimentally
\cere not accurate enough to predict adiabatic temperature increases. For design purposes, temperature increases are calculated from product distribution and heat of formation data. Catalyst in the reactor was supported and topped by a layer of Ij- to 3,:16-inch-dianieter inert alumina balls. Each reactor contained 300 cc. of catalyst. Since the goal of the study was to obtain data that would permit commercial design, no initial or transient high catalytic activity could be countenanced. It \cas necessary to ensure that the catalyst, when charged to the unit. had a n activity corresponding to the mid-point of its projected commercial life. To accomplish this, catalyst \cas prepared in a manner that \could simulate plant production and was then artificially aged by thermal treatment a t temperatures above those that \could be experienced in the process. N o activity decrease could ever be detected in the course of experimentation either because of time on-stream or successive regenerations. Mass spectrometer analyses were used to analyze liquid and gaseous streams. Toluene was used as the charge stock.
Table 111.
Detol Heavy Residue Analyses Analysis. I l t . $; 5070 97% CcHjCH;, CsHjCH3 cant'. conv. Biphenyl 35.6 83.1 Methylhiphenyls 13.8 ... Dimethylbiphenyls 3.5 , . . Biphenylmethane 2.1 ...
Methylbiphenylmethanes Bihenzyl Fluorene Others
Table IV.
l o w Coking Runs in Adiabatic Pilot Unit
Prpssurr.
P.S. I. A . 515
1200 1200
Bed Outlet Toluenp Temj. of Each Conwrszon, Reactor, a E: % 1203 53 8
Contact Tzme, Sec . 36
.Yo. of Reactors 2
70 71
3 3
1168 1174
142 144
4 1
1187 1200
Table V.
97.6 76.0
Typical Detol Yields
70-95'
conversion level
If%. of Fresh-Feed Aromatics
In later studies, charge stocks containing xylenes and substituted naphthalenes were evaluated in an analogous manner. Subsequent studies \cere also made to demonstrate the effects of impurities in the charge stocks, such as paraffins: naphthenes, olefins, sulfur compounds and nitrogen compounds. None of these were found to affect the catalyst activity. Heavy residues produced in the process \cere not vaporized into the mass spectrometer. Consequently, liquid product analyses were obtained on a residue-free basis. Residue quantity was determined by permitting a weighed sample of product to evaporate. Residue composition was determined by vapor fractometer analyses. Table I11 shows residue composition determined at tic0 conversion lexrels. At low conversions, the residue is mainly biphenyl and fluorene. At high conversions, biphenyl is the major product. T h e quantities of residue produced in the Detol process are so small that there is little economic incentive for recovery as a salable by-product. T o determine the purity of the benzene produced, liquid products were batch distilled a t 25 plates, 5 to 1 reflux ratio. The heart cut of benzene, representing 80 vol. % of the benzene in the sample, was analyzed for bromine index by ASTM method D 1492-57T ( I ) . The benzene was then clay treated in a conventional manner. Mass spectrometer analysis was used to determine purity. and thiophene was determined by ASTM method D 1685-591 (2). Adiabatic pilot unit runs were made a t operating conditions u i t h coking rates low enough to be practical for commercial operations and a t the same time covering a broad enough range of conversion levels to permit the establishment of a reaction model that would permit design optimization. In addition, runs \rere made a t conditions that were highly coking and which fouled the catalyst. This permitted catalyst activity and product distribution to be traced over a n entire period between regeneration. The coking information permitted establishment of a separate coking model, which was also incorporated into the computer program. Table I V lists typical runs made, the operating conditions, and the conversions obtained in the adiabatic experimental program. The yield data for both the adiabatic and the isothermal runs were correlated as a function of toluene conversion. LVithin experimental accuracy, hydrogen consumption and methane production are stoichiometric-1 mole per mole of toluene converted. No parameters of operating conditions were found. Toluene conversion was found to be an adequate representation of reaction severity. I n all of the work. coke production was negligible \\hen expressed as weight per cent of charge. Typical yields of net reactor products a t the 70 to 95Yc conversion level are summarized in Table \..
Charge Toluene
Consumed h \ droqen Total
Net products Benzene Heavy residurl Methane Ethan? Propane Butanes
C , and
+ (satulated)
T o t a1
Bt-nzene selecti\ i t \ .('mole
100.0 .2 _2_ 102.2 81.6 1.2 17.3 1.0 0.6 0.1 .4 _0_ 102.2 96.2
Discussion of Results
The pilot plant study permitted correlation of )ields and selectivities over a conversion range of 30 to 977,. Table \-I compares some predicted and observed conversions obtained by regressing the data for the low coking runs into the kinetic model. The activation energy corresponding to the lowest standard error ( ~ 3 . 9 7 is ~ )35.4 kcal. per gram-mole for the reaction over the unfouled catalyst. This can be compared with the value of 50 i 5 kcal. per gram-mole established by Silsb) and Sawyer (5)for the thermal reaction. The standard error for predicting conversion is lcithin the accuracy required for commercial design. VOL. 2
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167
Table VI.
Conversions Predicted by Computer Model for l o w Coking Runs
Predicted Cumulative Conversion, 7 0 Reactor Reactor Reactor Reactor 7 2 3 4
79.4 40.3 31.9 33.9
...
...
63.1 56.4 57.4
...
... ...
71.3 74.9
86.5
...
0hserved
conversion at Last Dzference, Reactor, 70 %
76.0 64.7 71.1 87.8
+3.4 -1.6 +0.2 -1.3
The data for the fouled catalyst runs were found to correspond to activation energies higher than 35.4 kcal. per grammole. The specific value was found to be a function of the coke level on catalysts. Figure 4 shows the Arrhenius plots for thermal operation and those established by the computer regression for the unfouled stabilized catalyst and for two levels of coking. Benzene Purity. Nitration grade benzene requires a minimum freeze point purity specification of 5.4” C. This specification thus allows no more than 0.1 vol. % impurities in the product benzene. Table VI1 lists benzene purities measured in three reactor adiabatic runs at various pressure levels. This table indicates that at the 1175” F. temperature level, benzene of 99.96+ vol. yo purity is invariably produced in the pressure range of 515 to 1015 p.s.i.a. The thiophene content of the benzene produced is shown in Table VI1 to be consistently in the range of 0.2 p.p.m. This thiophene level is characteristic of petroleum derived benzene. The bromine index of the benzene produced increased from 15.7 to 69.9 as pressure was raised from 515 to 1015 p.s.i.a. Bromine index is a measure of unsaturates: such as cyclohexene, present in the benzene. If not removed, the unsaturates will discolor H2S04, resulting in an undesirable acid wash color. Conventional clay treating is adequate to reduce the bromine index in any of the benzenes produced to 1 or less, which corresponds to an acid wash color of zero. The increase in bromine index from 15.7 to 69.9 is of little economic consequence, Clay life is generally in the range of 8000 barrels of benzene per ton of clay. Conclusions
By use of the kineric reaction model and the computer regression, it was possible to characterize the Detol process with a minimum of experimentation. By-product yields, benzene selectivities, coking rates, and benzene purities that were obtained in the adiabatic experiments were directly translatable to commercial design.
168
I & E C PROCESS D E S I G N A N D D E V E L O P M E N T
Table VII. Total pressure, p.s.i.a. Bedoutlettemp., O F , Purity (vol. %) Cg olefins Methylcyclopentane Cyclohexane Benzene Bromine index Thiophene, p.p.m.
Product Benzene Analyses
515 1178
715 1183
815 1178
1015 1167 0.01
0.02
99.98 15.7 0.08
0.01 99.99 24.8 0.10
0.03 0.01 99.96 52.0 0.22
0.03
99.96 69.9 0.22
In the resulting design model, the rate constant activation energy and frequency factor corresponding to any desired level of catalyst fouling are utilized. By-product yields a t any given conversion level establish the temperature rise and hydrogen consumption of the reactions. The design model is calculated on the computer in progressive conversion increments for as many reactors in series as are desired. The validity of the development approach was proved in the operation of the first commercial unit a t Houston, Tex. The conversions obtained by the commercial unit at varied operating conditions are essentially the same as the values predicted by the computer. Benzene selectivity was found to be 96 mole 7 0 and the purity of the benzene produced is 99.95 mole %. The thiophene content of the benzene from the commercial unit is 0.2 p,p,m., just as in the pilot plant. I n all respects, the commercial results have matched the pilot plant results. literature Cited
(1) Am. Soc. Testing Materials, Philadelphia, Pa., “1958 Book of ASTM Standards,“ Pt. 8, p. 1340. D1492-57T, 1959. (2) Zhid., “1959 Suppl.,” P t . 8, p. 196. D1685-59T, 1959. (3) \ , Kirk, R. E., Othmer, D. F., “Encyclopedia of Chemical Technology,” Vol. 5 , p. 145, Interscience, New York, 1950. (41 Matsui. H.. Amano. .4. Tokuhisa. H., Bull. Japan Petrol. ’ . Znst. 1, 67 (March 1959). (5) Silsby, R. I., Sawyer, E. W., J . Appl. Chem. (London) 6 , 347 1956). (6) Stijntjes, G. L. F., Voetter, H., Roelofsen, E. F., Verstappen, J. J., Erdoel Kohle 14,1011 (1961). (7) Towell, G. D., Martin, J. J., A.Z.Ch.E. J . 7, 693 (1961). (8) Tsuchiya, Hashimoto, .A., Tominaga, H., Masamune, s. Bull. J a p a n Petrol. Znst. 1,73 (March 1959). (9) Weiss, .4. H., Hydrocarbon Process. Petrol. Refiner 41, 185 (1962). (10) \Yeiss, A . H., Maerker, J . B., Sewirth, R., Oil Gas J . 60, 64 (Jan. 22, 1962). \
I
RECEIVED for review August 7, 1962 ACCEPTEDDecember 5. 1962 Division of Petroleum Chemistry, 142nd Meeting, ACS, Atlantic City, N. J., September 1962.
