Ind. Eng. Chem. Res . 1987,26 , 1895-1901 1895

Murray, D. W.; Clarke, C. T.; MacAlpine, G. A.; Wright, P. G., Paper

Ohkatsu, Y.; Kikkawa, K.; Osa, T. Bull. Chem. SOC. Jpn. 1978,51,

Pearson, R. G. J. Chem. Educ. 1968, 45, 581, 643. Rowe, C. N.; Fein, R. S.; Kreuz, K. L. Prepr.-Am. Chem. SOC., Diu.

Sexton, M. D. J. Chem. SOC., Perkin Trans. 2 1985, 59. Spedding, H., private communication, Esso Chemical Ltd., 1980. Warne, T. M.; Vienna, P. C. Lubr. Eng. 1984, 40, 211. Yoshida, T.; Inoue, K.; Watanabe, H. J . Jpn. Pet. Znst. 1984,27,519.

Receiued for reuiew December 10, 1986 Accepted June 2, 1987

821236, 1982; Society of Automotive Engineers, New York.

3606.

Pet. Chem. 1958, A121.

Effect of Sodium Impregnation on Catalyst Performance When Hydrotreating a Coal-Derived Liquid

James R. Baker, Robert L. McCormick, and Henry W. Haynes, Jr.* Department of Chemical Engineering, University of Wyoming, Laramie, Wyoming 82071

A commercial Ni-Mo/alumina catalyst was loaded with percentage quantities of sodium and tested for activity and activity maintenance while hydrotreating a coal-derived distillate under high severity conditions. The original catalyst exhibited a large high-temperature peak during the tempera- ture-programmed desorption of tert-butylamine. This peak, which is an indicator of catalyst acidity, was reduced and ultimately eliminated by impregnating the catalyst with increasing amounts of sodium. At 5 w t % Na20, no acid peak remained. The sodium loadings had little effect on activity or activity maintenance for hydrogen uptake or specific gravity reduction as compared to the original catalyst. The hydrodenitrogenation activity was reduced, however. Activity maintenance for both the original catalyst and the sodium-treated catalyst was excellent, so it was not possible to observe an effect of sodium addition on activity decline. However, the sodium-loaded catalyst was coked to a lesser extent than the original catalyst.

There is little doubt that the economics of direct coal liquefaction will someday become competitive with con- ventional sources of liquid fuels. When this day comes, the process chosen for commercialization will likely be the one which makes the most efficient use of heterogeneous catalysts. Catalysts are critical to any coal liquefaction process, but it is not yet established with certainty where the catalysts will be located. Recent trends away from single-stage processing (H-coal) to two-stage thermal- catalytic (ITSL) processing were motivated to a great ex- tent by the need to protect the catalyst. Most recently the trend has been toward a close couple configuration where catalyst is again added to the liquefaction reactor. Ex- perience at the Wilsonville facility has demonstrated that distillate yields of the order of 70 wt 90 (maf basis) are obtainable with Illinois No. 6 coal (Spencer, 1987). The effective utilization of catalysts in the refining of coal- derived liquids will also be of critical concern.

One of the major challenges in the development of a successful coal liquefaction process is to prepare catalysts with improved activity maintenance characteristics. There are at least four features of a coal liquefaction environment that may be responsible for the rapid catalyst deactivation that is universally observed. (1) The high severity con- ditions in the primary liquefaction reactor may result in sintering of the catalyst and rapid formation of coke. (2) The hydrogen-deficient nature of the feedstock may con- tribute to rapid coke formation. (3) The high molecular weight components of the feedstock resist hydrogenation and include basic nitrogen and hydroxy polycyclic hy- drocarbons known to the active coke precursors. (4) The presence of coal ash and organometallic compounds may lead to deactivation by a physical blockage of catalyst pores.

Recently we initiated a research program to explore the causes of catalyst deactivation in a coal liquefaction en-

* To whom correspondence should be addressed.

vironment and to take steps to develop improved catalysts as opportunities become apparent. A systematic approach to the problem dictates that we make an effort to identify the features of the coal liquefaction system referred to in the previous paragraph which are primarily responsible for the deactivation. In the present paper, we report the results of a high-severity hydrotreating study of a hydro- gen-deficient coal-derived distillate. Our feedstock con- tains no residuum material, nor does it contain coal ash. The factors potentially responsible for deactivation of the catalysts are thus limited to (1) and (2) above.

While our primary interest is in coal liquefaction where the principal function of the catalyst is to replenish hy- drogen donor, i.e., to hydrogenate, we will also report re- sults pertaining to heteroatom removal. This information may be of interest to those involved in developing catalysts for coal liquids refining or upgrading applications.

Most studies indicate that coking is a major cause of catalyst deactivation in coal liquefaction (Ocampo et al., 1978; Cable and Massoth, 1985; Stohl and Stephens, 1985). I t is known that coke can be formed by acid-catalyzed reactions involving carbonium ions and by thermal surface reactions involving free radicals (Trimm, 1983). Appleby et al. (1962) demonstrated that surface acidity plays the dominant role in coke formation over cracking catalysts. Lewis acidity has been correlated with coke formation (Scaroni et al., 1984), and the coke-forming mechanism is believed to involve an interaction of Lewis sites with the unpaired electrons of nitrogen bases. While Bronsted acidity is also associated with coke information (Tanabe, 1970), some researchers have proposed that Bronsted sites may actually decrease coking by converting coke precursors to volatile products (Brunn et al., 1976; Furimsky, 1982).

The present trend in coal liquefaction research is to use conventional catalyst preparations designed for heteroatom removal in heavy petroleum feeds. These catalysts are almost exclusively Ni-Mo or Co-Mo on alumina supports and, as reported by several researchers, are intrinsically

0888-5885/81/ 2626-1895$01.50/0 0 1987 American Chemical Society

1896

acidic (Segawa and Hall, 1982; Boorman et al., 1985; Kiviat and Petrakis, 1973). The implications in the literature suggest that it might be possible to reduce a catalysts' coking tendency by neutralizing the acid sites. One ap- proach to eliminating surface acidity in the Mo/A1203 system is by the addition of alkali metals (Ratnasamy et al., 1974). However, there are apparently conflicting re- ports of the effects that such a treatment may have on catalyst activity. Muralidhar et al. (1984) observed adverse effects of sodium addition on initial activities in studies with model compounds. At low levels of sodium (0.5 wt % Na,O), the hydrodesulfurization (HDS) activity was strongly depressed, but hydrogenation (HYD) activity was only marginally lowered. Higher loading levels (5.0 wt '70 NazO) resulted in a complete loss of HDS activity and a drastic reduction of HYD activity. Kovach et al. (1978) reported a severe loss of hydrogenation activity of a Co- Mo/A1203 catalyst used to hydrotreat coal liquids when percentage quantities of sodium were added. On the other hand, Kelly and Ternan (1979) found the specific activities (for pitch conversion and hydrodeoxygenation) of alkali- promoted Co-Mo/Al,03 catalysts to increase with addition of the metal when hydrotreating Athabasca bitumen. The addition of sodium resulted in a marginal improvement of HDS activity and a reduction in HDN activity. Of prime interest to our investigation is the observation that coke formation was reduced by the sodium promotion. Boorman et al. (1982) observed that sodium addition to a Co-Mo/Alz03 catalyst was generally detrimental when processing Athabasca bitumen but that the effect was less detrimental when the sodium was added during the last step of the preparation.

Unlike these earlier reports, which were essentially concerned with initial activity measurements, the present study focuses on the effects that sodium addition may have on steady-state activities after the initial period of carbon lay down. A primary objective was to determine to what extent, if any, sodium treatment might affect carbon for- mation. The acidity of the catalyst was also investigated.

Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987

Experimental Section

Acidity Characterization. While a complete de- scription of the acidic nature of a solid must involve the determination of acid site density, acid site strength, and whether the site is of the Lewis or Bronsted type, this study was only concerned with determining the point a t which surface acidity had been completely neutralized by sodium addition. Therefore, a Temperature Programmed De- sorption (TPD) method, similar to that described by Nelson et al. (1983) and Mieville and Meyers (19821, was used. The heart of this system is a modified Gow-Mac Model 550 gas chromatograph. As illustrated in Figure 1, a small amount of catalyst is contained in a sample holder that can be heated by the GC furnace. Helium gas at near atmospheric pressure flows over the sample. After pretreatment to clean the surface, tert-butylamine (TBA) is preadsorbed on the surface by a continuous injection of TBA into the carrier gas upstream of the sample. Once a steady-state output is achieved, as indicated by the thermal conductivity (TC) detector, TBA injection is stopped and the excess gas is flushed out. After the TC detector output returns to base line, the catalyst is heated to create a linear rise in temperature with time. A small thermocouple, inserted in the catalyst, measures the tem- perature, and the downstream TC detector measures the thermal conductivity change of the sample stream relative to the reference stream. By use of a high flow rate to minimize diffusion resistance and readsorption, the TC

TEMPERATURE PROGRAMMED

OVEN - - - -_ - I FC SAMPLE ' I TCD

HOLDER q-4-

BUBBLER M I X T U R ~

Figure 1. Temperature-programmed desorption apparatus for acid sites characterization.

Figure 2. Bench-scale hydrotreating apparatus.

detector response is nearly proportional to the rate of desorption.

As the catalyst is heated, TBA and various decompo- sition products are eluted. As the temperature increases, the desorption rate increases, eventually goes through a maximum, and drops back to zero as the surface is de- pleted of adsorbate. A desorption spectrum usually ex- hibits two overlapping peaks. The low-temperature maximum (a: peak) has been shown to be caused by de- sorption of physically adsorbed TBA. The second high- temperature maximum (@ peak) is then a result of de- sorption and/or decomposition of chemisorbed TBA. The area beneath the @ peak is an indicator of the number of acid sites.

Since it has been reported that sulfiding eliminates Bronsted acid sites on Mo/A1203 catalysts (Ratnasamy and Knozinger, 1978), the catalyst samples were presulfided prior to TPD analysis. Sulfiding was accomplished by a procedure similar to that recommended by Thomas (1970). Air is flushed from the sample by flowing nitrogen at am- bient temperature. A H2S/IHz gas mixture (9.9% H,S) is then passed over the sample while the temperature is raised to 400 "C over the course of an hour and held at 400 "C for an additional hour. After cooling back to ambient, the H,S/H, flow is stopped and the excess sulfiding gas is flushed from the sample with flowing nitrogen.

Bench-Scale Hydrotreating. A schematic diagram of the trickle bed reactor is presented in Figure 2. For deactivation studies, a charge of approximately 4 g of catalyst, diluted 1-3 times with quartz chips (10-30 mesh) was used. The feedstock was a Pittsburg coal tar fraction obtained from US. Steel, for which physical properties are presented in Table I. Pure hydrogen and liquid feed are passed over the fixed bed of catalyst in a cocurrent downflow mode. The liquid product is collected in high- pressure accumulators. Gas from the high-pressure accu- mulators leaves through a back-pressure regulator, passes through the wet test meter, and is collected in a butyl

Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987 1897

Table I. Analyses for Pittsburgh Coal Tar Fraction (F-4 Feedstock)

w t % w t %

C 89.12 350- O F 0 H 6.77 350-500 O F 15.0 N 0.72 500-650 O F 82.5 S 1.42 residue 2.5 0 (diff) 1.97 sp. grav. (60/60 O F ) 1.0716

Table 11. Properties of Fresh. Calcined Catalysts" Na-impregna-

ted Shell Shell 324111 324M

~

0 5% Na,O on Shell 324M . Sulfide

I O % Na,O on Shell 324M-Su l f ide

BET surface area, m/g pore vol, cm3/g median pore diameter, 8, NiO, wt % Moo3, wt 70

Na20, wt 96 P*O6, wt %

A1203, wt %

157 (146) 0.49 (0.46) 86 (100) 3.5 19.4 62.4 0.1 6.9

133 0.44 88 (100) 3.2 17.6 58.4 4.7 6.5

50% Na,O on Shell 324M-Sulf lde

a Data graciously provided by the Amoco Oil Company. Values in ( ) were determined in house.

rubber gas bag for analysis. The unit is operated contin- uously, but the data used for balance periods are typically collected over a period of 4 h.

Prior to the start of a run, the catalyst was presulfided by a procedure similar to that reported by Hallie (1982) to provide the optimum activity for Ni-Mo catalysts. With a reactor pressure of 1.7 MPa (250 psia) and a hydrogen flow rate of 2.5 L(STP)/h, a sulfiding feed of 5 wt % CS2 in cyclohexane was passed over the catalyst bed at a rate of 10 g/h. The reactor temperature was then raised from ambient to 200 OC at a rate of 60 OC/h. After the tem- perature was held at 200 OC for 2 h, it was again raised at a rate of 60 OC/h to 300 "C, held for an additional hour, and then allowed to cool back to ambient temperature under flowing hydrogen and sulfiding feed.

During a hydrotreating run, the liquid feed was supplied by a Ruska Model 2252-BI proportioning pump. The hydrogen flow rate was set by using a Brooks Instrument Model 5871-A mass flow controller. Every effort was made to maintain the reactor isothermally. This necessitated the fabrication of a jacket surrounding the reactor which allowed air to be injected to cool hot spots in the bed. With this arrangement, temperature profiles with 4-6 OC vari- ations were typically observed.

Hydrotreating Catalyst. The catalyst used in this study was a commercial Ni-Mo/A120, preparation (She11 324M) which has seen service at the coal liquefaction pilot plant facility in Wilsonville, AL. The sodium-modified Shell 324M catalysts were prepared by the incipient wetness technique using freshly calcined Shell 324M cat- alyst and an aqueous solution of sodium nitrite a t the desired concentration. After impregnation, the catalysts were calcined at 400 "C for 2 h in flowing air. Properties of both catalysts are presented in Table 11. A notable feature of the Shell 324M catalyst is its unimodal pore structure. I t should also be noted that the "as received" catalyst contains a small quantity of sodium (0.1 wt % Na20).

Results The original and sodium-impregnated catalysts were

first subjected to TPD analysis for acid sites characteri- zation. Samples of the catalysts were then tested in the bench-scale hydrotreating unit for extended time periods in order to observe the activity decline.

90 z I30 170 210 250 290 330 '0

TEMPERATURE, "C Figure 3. TPD results for sodium-loaded Shell 324M.

TPD Analyses. TPD analysis of the Shell 324M cat- alyst reveals a large well-defined high-temperature peak (p peak). This is typical of alumina-supported molybde- num catalysts that we have studied, and the acidity ap- pears to be associated with the incorporation of molyb- denum into the catalyst. A y-alumina support free from catalytic metals typically exhibits a very broad TPD spectrum which does not return to base line until near the end of the temperature program. When molybdenum is added to the support, a sharp p peak is superimposed on the TPD curve.

The effect of sodium loading on the p peak is clearly evident in Figure 3. This peak is completely extinguished at the 5.0% NazO level. Note that all the samples have been sulfided. An electron microprobe analysis of the 5% loaded catalyst indicated that the sodium was uniformly distributed across the diameter of the pellet. (This in- formation was graceously provided by Dr. M. A. Pacheco of the Amoco Oil Company.)

Computation of the Reaction Kinetics. Because of contacting inefficiencies characteristic of laboratory trickle bed reactors of the type employed in this investigation (Satterfield, 1975), it was not possible to develop reliable kinetics that might be used for process design purposes. However, it was necessary to develop the kinetics in suf- ficient detail that corrections for small day-to-day devia- tions from target conditions could be made. Also, the most convenient and meaningful way to follow catalyst deac- tivation is to plot a rate constant. Accordingly, crude kinetics models were developed for the reactions of primary interest to this study-hydrogenation and the related specific gravity reduction.

Catalyst HYD deactivation was monitored by following a HYD rate constant. For hydrogen content, a kinetics model was selected based upon the assumptions of a plug flow reactor and a first-order reversible reaction. The integrated kinetics expression takes the form

where k;I is the observed first-order rate constant and rW is the weight hourly space time. The H s are weight per-

1898 Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987

in

in h

- 9 ‘ c 3 3 8 2 c ‘ 8

p z c 9 : I ~ ?rz;-’

Figure 4. Parity plot for HYD kinetics model.

centages of hydrogen and the subscripts F, P, and E refer respectively to the feed, product, and equilibrium hydrogen contents.

To correct for small deviations from the target reactor temperature (TT), the observed rate constant was corrected in accordance with the Arrhenius law:

y1

yi

S P E C I F I C GRRVITY 60’60

Figure 5. gravity of products derived from F-4 feedstock.

Table 111. Thermal Experiment Results

Correlation between hydrogen content and specific

The term T E is an effective isothermal temperature ob- tained by integrating over the experimental temperture profile according to

1 L exp(-E,/RTE) = -s exp(-E,/RT) dz (3) L o

where L is the catalyst bed length and Tis the temperature a t any distance z into the catalyst bed. The discrete ex- perimental temperature profiles were fit with a cubic spline approximation, and the integral was evaluated numerically by using Simpson’s Three-Eights rule.

For this method to be used, a value for both E, and H E must be determined. This was accomplished by perform- ing a hydrotreating r w devoted entirely to developing data suitable for kinetics modeling. Over the course of 367 h, 16 balance periods were performed at various combinations of three target weight hourly space times ( T ~ = 0.46,0.92, and 1.84 h) and three target effective isothermal temper- atures ( T E = 385, 415, and 440 “C). The collected data were regressed on a rearranged form of eq 1

where Hp,c is the calculated hydrogen content of the liquid product and the observed rate constant is separated into an intrinsic rate constant (h’k) and a deactivation factor (OH). The quantity HE is assumed to be linearly related to TE, in degrees kelvin, Le.,

(5) and k ’;I has an Arrhenius temperature dependence

krrH = b, exp(E,/RTE) (6) It is further assumed that the term OH is related to the cumulative oil/catalyst weight ratio (x) according to

( 7 ) A deactivation factor of unity thus corresponds to the HYD activity when an arbitrary 20 g of feed has been processed for each gram of catalyst.

By use of eq 4-7 in conjunction with a Taylor Series Linerization procedure, optimum (in the least-squares sense) values of the model parameters were found. Details

HP,C = HE + (HF - H E ) exp(-OHk”Hrw) (4)

HE = blTE + b2

8, = 1 - b,(x - 20)

BMCO6-1 feed thermal product product

c, wt % 89.12 89.31 89.37 H,wt % 6.77 7.00 10.18 s, wt % 0.72 0.90 tr N, wt % 1.42 1.30 0.024 0, wt % (diff.) 1.97 1.49 sp. grav. (60/60 “F) 1.0716 1.0613 0.9346

of this procedure were fully described elsewhere (Baker, 1986). This model describes the hydrogen content data quite well as evidenced by the parity plot of Figure 4. The observed values are those determined by analysis of the liquid yield period products, while the calculated values are those predicted by eq 4-7 using the optimum values of the model parameters.

A similar workup was performed on the specific gravity data. An identical model was proposed with specific gravity replacing hydrogen content. Again, the model describes the observed specific gravities quite well. This similarity in the hydrogenation and specific gravity re- duction results was no surprise since for catalytic hydro- treating studies on a given feedstock it is often found that the hydrogen content of the product has a high correlation with the product specific gravity. As Figure 5 illustrates, this was indeed true for our study.

In addition to the process variable study, a short thermal study was performed to determine the magnitude of any thermal reactions. The operating conditions were identical with those of the deactivation runs; T E = 427 “C (800 O F ) ,

P = 13.9 MPa (2000 psig), and a liquid feed rate of 8.77 g/h. A comparison of the thermal product to the feed and to a typical catalytic product indicated that thermal re- actions are relatively insignificant, Table 111.

Deactivation Studies. Four deactivation runs were completed a t the nominal conditions of 13.9 MPa (2000 psig), 427 “C (800 O F ) , and a weight hourly space time of 0.46 h. The hydrogen treat rate was maintained at 0.63 L(STP)/cm3 (3500 SCF/BBL). The catalyst description and duration of each run are summarized in Table IV. Operation was generally good, although some plugging was experienced due to iron deposits forming near the bed entrance. This necessitated an early termination of run BMCO6. A power outage forced a premature termination of run BMC07. Runs BMC08 and BMCOS were termi- nated voluntarily after reaching the targeted 400+ h of operation.

In Figures 6 and 7 the corrected rate constants, based respectively on hydrogen content and specific gravity, are plotted as a function of the amount of feed processed by

Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987 1899

? -

Table IV. Deactivation Runs (Nominal 2000 psi, 800 OF, 2 g/h/n)

run duration wt of

feed/& of run catalvst hours catalvst

~~ ~ ~

BMCO6 Shell 324M 172 374 BMC07 3.02% Na20/Shell 324M 242 522 BMC08 Shell 324M 409 891 BMCOS 4.97% Na20/Shell 324M 410 896

0 - BNC06 0 - BMCOB

0

CUMULRTIVE WT. FEE0 / WT. CRTALYST

Figure 6. Hydrogen uptake deactivation curve. Shell 324M cata- lyst.

9(: 0 - 8MC06

0 - BMCOB

0 IO0 200 300 400 500 600 700 800 900 CUMULRTIVE WT. FEEO / WT. CRTRLYST

Figure 7. Specific gravity reduction deactivation curve. Shell 324M catalyst.

the unmodified Shell 324M catalyst. Surprisingly, these HYD deactivation curves reveal that little activity loss has occurred even after the catalyst has been on stream for a relatively long period of time. Recently Cillo et al. (1985) have reported that in processing a 50-50 mixture of hy- drogenated creosote oil and residuum, the Shell 324M catalyst was able to maintain a high HYD activity for a period much longer than that used in this study (ap- proximately 2000 g of feed/g of catalyst).

In Figures 8 and 9 the HYD deactivation curves for the sodium-modified catalysts are compared with the results for the original catalyst (dashed line). These results in- dicate that sodium addition does indeed lower the initial HYD activity, though not to the degree frequently claimed in the literature. The long-term activity, however, is relatively unaffected by sodium addition. It may well be, as suggested by Boorman et al. (1984), that when sodium is added after metals impregnation and calcining, little

0 100 200 500 400 500 600 700 800 900 CUNULRTIVE WT. FEEO / WT. CATALYST

Figure 8. Hydrogen uptake deactivation curves. Sodium-impreg- nated catalysts.

YI

0 100 200 300 400 500 600 700 800 CUMULATIVE WT. FEED / UT. CATALYST

3

Figure 9. Specific gravity reduction deactivation curves. Sodium- impregnated catalysts.

4 BYCOB 4 B M C O 9

I I I I 200.0 400.0 600.0 800 .0 loo(

CUM. v i . OIL PER w i . OF CATALYST

0

Figure 10. Hydrodenitrogenation deactivation curves for the ori- ginal and sodium-modified catalysts.

effect is observed on the HYD activity. In any case, the addition of sodium to the Shell 324M catalyst does not seem to have caused any great change in the overall HYD performance.

While hydrogenation is the main function required of a coal liquefaction catalyst, heteroatom removal is im- portant to upgrading catalysts. Hydrodesulfurization at the conditions of this study was almost complete, and it was not possible therefore to obtain definitive HDS results.

1900 Ind. Eng. Chem. Res., Vol. 26, No. 9, 1987

Table V. ProDerties of Sulfided Catalysts, Fresh and Spent Na-impregnated

Shell 324M Shell 324M fresh fresh

sulfided spent sulfided spent BET surface area, m2/g 121 123 132 137 pore vol (>60-8, dia.), cm3/g 0.35 0.19 0.35 0.19

wt 70 change on combustion 16 8 Na/Al wt ratio 0.17 0.10 re1 acid density, m-2 0.091 0.051 neg. neg.

However, it was possible to observe hydrodenitrogenation during the deactivation runs. First-order rate constants for HDN are plotted in Figure 10 for runs BMC08 and BMCO9. In some cases small corrections to the target temperature were required. An activation energy of 14.8 kcal/gmol was assumed for this purpose in accord with the results of Ternan and Brown (1982).

The most obvious trend from these results is that the data for run BMCO9 lie consistently below the data from run BMC08. Sodium clearly serves as a poison for the HDN activity. Evidently, the acid function of the catalyst interacts preferentially with basic nitrogen compounds, and elimination of this acidity serves to reduce the HDN se-' lectivity.

A second trend, namely the increasing HDN activity with time on stream, is less obvious, as the data are scattered. Some increase in activity during the early part of the run might be attributable to an incomplete pre- sulfiding of the catalyst. The maximum temperature during presulfiding was only 300 "C (572 OF). While the procedure falls within the guidelines recommended by Hallie (1982), some adjustment in the procedure may be warranted in future studies. In particular it may be de- sirable to raise the final temperature to the reaction tem- perature.

Characterization data for the fresh sulfided and spent catalysts are compared in Table V. The spent catalysts were Soxhlet extracted in pyridine prior to analysis. The pore volumes in this table were calculated from mercury porosimetry data and therefore reflect volume in pores > 60-A diameter. There is clearly a substantial loss in pore volume in the spent catalysts. This is accompanied by a decrease in the average pore diameter. Surprisingly the surface areas remain constant. The weight percent change on combustion may be taken as an indication of carbon content. This value is substantially lower for the sodium- impregnated catalyst than for the original catalyst. It appears therefore that impregnating the catalyst with so- dium may serve to reduce coking tendency. Another in- teresting number is the Na/Al weight ratio from analysis of the fresh and spent sodium-treated catalysts. It appears that sodium is removed from the catalyst during the course of the run. This may provide an alternative explanation for the apparent increase in HDN activity observed in Figure 10. The last entry in Table V, the "relative acid density", is calculated from the ratio of the P-peak area to the sample surface area. The number by itself is of little utility; however, it does provide a means for comparing relative acidities among a group of catalysts. The original Shell catalyst appears to experience some loss of acidity during processing. Control sample testing with the fresh catalyst indicated that this was not a consequence of the pyridine extraction procedure.

Conclusions The addition of percentage quantities of sodium to a

commercial Ni-Mo/A1203 hydrotreating catalyst pro-

median pore diameter, A 100 75 100 75

gressively eliminates acid sites as determined by the TPD of an organic base. The introduction of sodium into the catalyst results in a reduction in initial activity for hy- drogenation and specific gravity reduction. However, the effect on "lined-out" activity is minimal. A key factor may be that the sodium is added after the catalytic metals. The HDN activity is suppressed by sodium addition. Both the original and sodium-promoted catalysts performed re- markably well, as little deactivation was observed after the early hours of the run. Evidently the high severity con- ditions employed and the hydrogen-deficient nature of the feedstock are not in themselves sufficient to deactivate the catalyst to the extent observed in coal liquefaction ap- plications. Work presently under way will address the influence that residuum material may have upon the deactivation. While it was not possible to observe an effect of sodium promotion on activity maintenance, sodium addition appears to retard the formation of carbon on catalyst.

Acknowledgment

This work is jointly supported by the U.S. Department of Energy (Grant DE-FG22-84PC70812) and the Amoco Oil Company. We are grateful for valuable consultations with Drs. M. M. Schwartz and M. A. Pacheco of Amoco and Dr. R. E. Tischer of the U.S.D.O.E.

Nomenclature b = regression parameter E, = activation energy H = weight percent hydrogen k = first-order rate constant corrected to target conditions k' = observed first-order rate constant k" = intrinsic first-order rate constant L = bed length P = pressure R = gas constant T = absolute temperature x = cumulative weights oil/weight catalyst Greek Symbols 0 = deactivation factor T~~ = weight hourly space time Subscripts C = calculated value E = equilibrium value (or effective isothermal value) F = feed value H = hydrogenation P = product value T = target value

Regis t ry No. Ni, 7440-02-0; Mo, 7439-98-7; Na, 7440-23-5.

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Received for review December 29, 1986 Revised manuscript received May 27, 1987

Accepted June 13, 1987

Conversion of CH4 into C2H2 and C2H4 by the Chlorine-Catalyzed Oxidative-Pyrolysis (CCOP) Process. 1. Oxidative Pyrolysis of CH3Cl

A. Granada, S. B. Karra, and S. M. Senkan" Depar tment of Chemical Engineering, Illinois Ins t i tu te of Technology, Chicago, Illinois 60616

The oxidative pyrolysis of CH3C1, representing the second stage in the chlorine-catalyzed oxidative pyrolysis (CCOP) of CH4, was studied in a flow reactor at about 980 "C and 0.68 atm. The presence of oxygen in the system decreased the extent of formation of carbonaceous deposits considerably without the formation of destructive flames. Under the reaction conditions studied, the formation of C2H2 and CzH4 with combined yields as high as 60% was observed, a t about 30% conversion of CH3C1.

Conversion of methane into higher molecular weight hydrocarbons is of immense practical significance. Methane is available in large quantities in natural gas, thus constituting an important raw material for the synthesis of higher molecular weight hydrocarbons. Processes exist to convert methane into acetylene, ethylene, and hydrogen by using high-temperature pyrolysis. However, at the high temperatures needed for the thermal decomposition of methane, the yields of more valuable liquid and gaseous products are too low due to the formation of excessive amounts of carbonaceous deposits (see for example Back and Back (1983) and references therein).

In an earlier patent, Gorin (1943) described a two-step chlorine-catalyzed process for the polymerization of methane. According to this process, methane is chlorinated first, forming chlorinated methanes (CM), followed by the pyrolysis of CM and formation of C2 and C2+ hydro- carbons and HC1 in the second step. The HC1 produced can either be converted into chlorine via the well-known Deacon reaction and recycled or can be used to oxy- chlorinate methane to form CMs, thus completing the catalytic cycle for chlorine. Recently, Benson (1980) patented a theoretical single-step process which involves the ignition of C12-CH4 mixtures and formation of flames. Later, Weissman and Benson (1984) studied the kinetics of nonflame pyrolysis of CH3CI.

As shown in these studies, although the decomposition temperatures of chlorinated methanes are considerably lower than those for methane, thus the destruction rates of useful products are lower, the formation of carbonaceous deposits, which include high molecular weight low vapor pressure hydrocarbons, tars, carbon, and soot, is still a

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problem (Gorin, 1943; Weissman and Benson, 1984). This renders the direct pyrolysis of CMs unattractive for practical applications.

I t must be recognized that although the formation of carbonaceous deposits may represent a minor route, it has major consequences. Because of the accumulative nature of the deposits, the reactor and transportation lines ulti- mately plug up, and this results in major process ineffi- ciencies.

In this paper, we report on the results of experimental studies of the oxidative pyrolysis of CH3C1, representing the second stage in the chlorine-catalyzed oxidative py- rolysis (CCOP) of CHI developed recently (Senkan, 1987). Product distributions measured both in the presence and absence of oxygen clearly show that O2 effectively and efficiently ameliorates the problem of formation of high molecular weight carbonaceous deposits and allows the production of C2H2 and C2H4 with high yields. A plausible reaction mechanism for the oxidative pyrolysis of CH3C1 is also discussed.

Experimental Section The experimental facility used is shown in Figure 1.

The reactor was a 2.1-cm-i.d. by 100-cm-long quartz tube, of which about 60 cm was placed in a three-zone Lindbergh furnace. The first zone of the furnace, which was about 15 cm long, was used to preheat the argon carrier gas. Small amounts of CH3Cl and CH3C1/ O2 mixtures were introduced into the preheated argon carrier gas by using an air-cooled probe through radially directed injection holes. The amounts of gases injected were deliberately kept small to ensure the rapid heat-up of the reactants and to preserve the near-isothermal conditions during the ex- periments.

0 1987 American Chemical Society