146
Ind. Eng. Chem. Res. 1992,31, 146-155
Catalytic Pyrolysis of Naphtha Biswadip Basu and Deepak Kunzru* Department of Chemical Engineering, Indian Institute of Technology, Kanpur 208016, India Steam pyrolysis of naphtha has been studied over a calcium aluminate (12Cao-7fil2o3) catalyst in a fixed bed reactor at atmospheric pressure in the temperature range 973-1123 K. Conversion of naphtha increased appreciably in the presence of the catalyst, reducing the temperature required for a given conversion by approximately 50 K. Compared to thermal pyrolysis, addition of the catalyst significantly increased the production of methane, ethylene, and propylene. At higher temperatures and relatively longer residence times, the formation of carbon oxides was increased in the presence of the catalyst. The concept of equivalent reactor volume has been used to estimate the rate constants for thermal, as well as catalytic, pyrolysis of naphtha. The activation energy for catalytic pyrolysis (124300 kJ/kmol) was found to be significantly lower than for thermal pyrolysis (217 500 kJ/kmol).
Introduction Naphtha and light hydrocarbon gases are the usual feedstocks for olefin production by pyrolysis. During pyrolysis, ethylene is the major product together with significant amounts of methane, propylene, butenes, and 1,3-butadiene. In addition, hydrogen, ethane, and pyrolysis oil are also produced in lesser amounts. The product yields depend on various factors, such as residence time, temperature, and average hydrocarbon partial pressure. High olefin yields are favored at high temperature, low hydrocarbon partial pressure, and high residence times. Several investigators have studied the kinetics and product yields during pyrolysis of naphtha (Kumar and Kunzru, 1985; Sahu and Kunzru, 1988; Bajus et al., 1980; van Damme et al., 1981). Recently, there has been interest in developing catalysts for naphtha pyrolysis, with the aim of reducing the required temperature for a given converstion as well as reducing the coke formation during naphtha pyrolysis. The published information on catalytic pyrolysis is very limited. Various types of catalysts have been tested with feedstocks, ranging from light aliphatic hydrmbons to heavy gas oils, with conflicting results. Fischer et al. (1974) used zirconium oxide deposited on alumina as a catalyst for hydrocarbon pyrolysis. Small amounts of alkali-metal oxides were incorporated in the catalyst, resulting in lower coke yields. Kolombos et al. (1977a,b) tested zirconium or titanium oxide as carrier and manganese(IV) oxide as the active component for pyrolysis of a petroleum distillate and heavy residue. Good yields of light olefins were reported, and no coke deposition was measured even after 2 h. Egiazarov et al. (1978) used indium oxide on pumice as a catalyst and reported high yields of ethylene and limited coke deposition. Tomita et al. (1977) have developed calcium aluminate catalysts for the pyrolysis of various hydrocarbons ranging from light naphtha to crude oil. The most promising catalyst consisted of 51.46% CaO and 47.73% A 1 2 0 3 , the main crystalline phase being 12Ca0-7A1203. With such a catalyst, for steam pyrolysis of naphtha at 1093 K, the ethylene yield was 39.8 w t %, whereas no coke deposition was observed even after 120 h of continuous operation. A disadvantage of the catalyst is the high yield of carbon oxides. Kikuchi et al. (1985) have studied the effect of temperature, pressure, contact time, and steam/oil ratio during the pyrolysis of naphtha over such a catalyst. In the temperature range of 1023-1173 K, CzH4and CH, yields increased monotonically with temperature, whereas
* To whom correspondence should be addressed.
the C3H, and C02yields passed through a maximum. The decrease in C02 yield at higher temperatures was attributed to the deactivation of the catalyst due to coking. Ellig et al. (1985) studied the effect of calcium oxide on the pyrolysis of n-heptane and aromatic hydrocarbons. Calcium oxide significantly increased the rate of pyrolysis of benzene, toluene, and 1-methylnaphthalene, whereas the rate of heptane cracking was affected only moderately. Significant yields of CO and COBwere observed when any of these hydrocarbons were reacted over CaO. Lemonidou and Vasalos (1989) tested various catalysts for the steam pyrolysis of n-hexane. The best results were obtained with 12Ca0-7A1203catalyst. Calcium aluminate treated with hydrogen at high temperature did not show any catalytic activity. On the basis of these results, they suggested that the catalytic activity of CaO is presumably due to the excess oxygen present in the crystal phase (Ca12A114033) of the 12Ca0-7A1203catalyst. Very limited information is available on catalytic pyrolysis of naphtha. The aim of this study was to inveatigate the effect of temperature, steam dilution ratio, and space time on the product yields during pyrolysis of naphtha over a calcium aluminate catalyst. Another objective was to determine the kinetics of catalytic pyrolysis of naphtha.
Experimental Section Catalyst Preparation. The calcium aluminate catalyst was prepared by mixing the required amounts of CaO and A1203 in the presence of a binder, followed by subsequent aging,dehydration and, finally, sintering. Thus, to prepare a 40-g batch of catalyst, 20.5 g of CaO and 19.5 g of calcined A 1 2 0 3 were thoroughly mixed with 20 mL of 1% aqueous solution of poly(viny1 alcohol), which was used as the binder. The mixture was molded into cylindrical pellets of 12-mm diameter and aged in an oven at 353 K for 24 h. After aging, the pellets were placed in a furnace, and the furnace temperature was increased at a rate of 200 K/h to 1773 K. During this heating, the pellets were dehydrated and the organic binder burnt off. The pellets were kept at 1773 K for 12 h to ensure complete sintering of the calcium oxide and alumina powders. The catalyst prepared using this procedure showed good crystallinity, as measured by X-ray diffraction (XRD), and had a high crushing strength. For use in the pyrolysis studies, the pellets were crushed and then sieved. Catalyst Characterization. The prepared calcium aluminate catalyst was characterized using X-ray diffraction and surface area measurements. The XRD was obtained on a Reich Seifert Is0 Debye Flex 2002 unit using monochromated Cu Ka radiation.
0888-5885/92/2631-0146$03.OO/0 0 1992 American Chemical Society
Ind. Eng. Chem. Res., Vol. 31, No. 1,1992 147
1 2
allla
U
4 II
15
7
II 1
1 Burette filled wlth water 2 Burette filled with naphtha 3 Pumps 4 U - tube manometer 5 Vaporizer 6 Preheater 7 Heating tape 8 Reactor furnace 9 Reactor tube 10 Circular plate for supporting catalyst 11 Thermocouple to control Ier 12 Thermocouple 13 Rotameter 14 N2 cylinder with regulator 15 Condensor 16 Separating funnel 17 Gas sampling valve 18 Gas flow meter
Figure 1. Schematic of the experimental setup.
The calculated interplanar distances were compared with the standard values. The diffraction pattern of the prepared catalyst confirmed that both 12Ca0-7A1203 and 3Ca0-A1203 crystalline phases were present, although 3Ca0-A1203 was present in minor amounts only. The total surface area of the freshly sintered catalyst was measured using the dynamic pulsing technique on a Micromeritics Pulse Chemisorb 2700 unit. Nitrogen was used as the adsorbent. The total surface area of the fresh catalyst was 3.2 m2/g, which is in good agreement with the value of 3.05 m2/g reported by Lemonidou and Vasalos (1989) for a calcium aluminate (12Ca0-7A1203) catalyst prepared at 1573 K. Apparatus and Procedure. Both the thermal and catalytic pyrolysis runs were conducted in a tubular reactor. A schematic diagram of the experimental setup is shown in Figure 1. Naphtha and water were taken in separate burets and pumped by reciprocating pumps to a vaporizer. The flow rate of water could be varied in the range from 0.2 to 1.8 mL/min, whereas that of naphtha, from 1.5 to 30 mL/min. In the vaporizer, naphtha and water were separately vaporized and mixed at the outlet of the vaporizer. This mixed stream was then passed through a preheater where it was heated to approximately 773 K. This heated stream then entered the reactor. To minimize heat losses, the line connecting the preheater to the reactor was heated. The tubular reactor was constructed from a 800-mm length of Inconel tube (i.d., 19 mm; o.d., 25 mm). The heated length of the reactor was 600 111111. The axial temperature profile in the reactor was measured with chromel-alumel thermocouples using two thermowells (o.d., 6 mm). One thermowell was attached to the top of the reactor, whereas the other was inserted from the reactor outlet. A circular plate (diameter, 19 mm), containing ten l-mm holes, was welded to the upper thermowell and served as the catalyst-retaining support. The reactor was heated in a two-zone furnace. The heating to the upper zone was controlled through a PI controller (Model 4014, Indotherm Instruments, Bombay, India), which was connected to the reactor thermocouple. The heating of the lower zone was controlled manually through a Variac. A metered flow of nitrogen was mixed with the reactor effluent at the reactor exit. This was necessary because nitrogen was used as an internal standard in the subse-
quent chromatographic analysis. The reactor effluent, together with the added nitrogen, passed through two condensors, connected in series. Most of the unconverted naphtha, liquid products, and steam condensed in the first condensor. A glycerol-water mixture was used as the coolant and circulated through the condensors using a refrigerated circulator (Model F-ZOHC, Julabo, Seelbach, Germany). The noncondensablea passed through a sample valve to a gas flow meter and were then vented. During the course of a run, the total condensed liquid, together with the volume and composition of the product gases, was measured at regular intervals. Initially, the composition and product flow rates varied with the run time but achieved a steady state in approximately 30 min. The total run time was varied from 2 to 6 h. At the completion of the run,the reactor was flushed with steam for 1h, and then either the reactor was decoked with air or the coked catalyst was removed from the reactor for subsequent analysis. The noncondensable gases, which included C1-C4 hydrocarbons, carbon monoxide, and carbon dioxide, were analyzed by gas chromatography using three columns, viz., Duropak, Porapak Q, and Carbosphere. The separation of N2,CH,, CO, and C02was carried out on a Carbosphere column using a thermal conductivity detector with H2 as the carrier. Determination of C1-CI hydrocarbons was done on a 3-m-length Duropak column using a flame ionization detector. Since ethane and ethylene came as a single peak on the Duropak column, these were separated on a Poropak Q column. The hydrogen in the product gases and the liquid products were not analyzed in this study.
Results and Discussion To study the effect of the 12Ca0-7A1203 catalyst on naphtha pyrolysis, runs with and without catalyst were conducted at atmospheric pressure and a temperature range of 973-1123 K. The straight-run and desulphurized naphtha feeds were obtained from IC1 Ltd., Kanpur, India. The boiling range of both naphthas was 313-433 K, the density, 700 kg/m3, and the average molecular weight, 90. The ASTM distillation analysis of the two naphthas was the same and is given in Table I. T h e sulfur content of the straight-run naphtha was 80 ppm, whereas that of desulfurized naphtha was less than 5 ppm. The range of
148 Ind. Eng. Chem. Res., Vol. 31, No. 1, 1992 Table I. Characteristics of NaDhtha Feed density (298 K) 700 kg/m3
j
e
r
a
l
l gas yield ,
ASTM Distillation temp, K 313 340 361 375 388 410 433
vol % distilled" IBP 10 30 50 70 90 FBP average molecular weight
-,,-40-
PI .>
90
ethylene v
"IBP = initial boiling point; FBP = final boiling point. 20 -
Table 11. Operating Conditions for Main Pyrolysis Runs parameters value desulfurized naphtha flow rate, g/min 1.18-2.67 steam flow rate, g/min 0.97-1.58 reaction temperature, K 973-1123 space time ( T ) , s 0.13-0.3 0.48-1.35 steam to naphtha weight ratio (6), kg/kg catalyst weight, g 1.5-5.0 3.0 f 0.5 catalyst size, mm
propylene
-
-
ethane
Figure 2. Product distribution vs W / m Nfor the steam cracking of desulfurized naphtha: A,overall gas yield; 0,ethylene; 0 , methane; A, propylene; 0 , ethane.
operating conditions for the main pyrolysis runs is shown in Table 11. Addition of the calcium aluminate catalyst significantly increased the overall gas yield, as well as the yields of the major gaseous products, such as methane, ethylene, and propylene. The effect of varying the ratio of the mass of catalyst to the mass flow rate of naphtha ( W / m Nfrom ) 54 to 178 ((kg of catalyst) s)/(kg of naphtha) at 1073 K is shown in Figure 2. In this set of runs, the flow rates of desulfurized naphtha and steam were kept fixed and W/mNchanged by varying the catalyst weight from 1.5 to 5.0 g. For the sake of comparison, a pyrolysis run with no catalyst ( W / m N= 0)is also shown on this figure. As can be seen from this figure, in the range of W / m Nfrom 0 to 107 ((kg of catalyst) s)/(kg of naphtha), yields of total gaseous product, methane, ethylene, and propylene increase significantly with an increase in W / m N however ; increasing W / m Nabove 107 ((kg of catalyst) s)/(kg of naphtha) (corresponding to a catalyst weight of 3 g) does not have a significant effect on the product yields. Hence, for all the subsequent experiments, the catalytic pyrolysis runs were made with 3 g of catalyst. It should be mentioned that although the fixed bed reactor had significant
axial temperature gradients, there was an isothermal zone in the central portion of the reactor where the catalyst was placed. Thus, for all the runs shown in Figure 2, the catalyst surface temperature was the same. The analysis of the experimental data in any pyrolysis study is rendered difficult by the inevitable temperature profile which exists in the reactor. The data are usually analyzed using a pseudoisothermal approach based on the equivalent reactor volume concept (van Damme et al., 1981; Kumar and Kunzru, 1985). The space times, given in Table I1 and Figure 2, were evaluated from the calculated equivalent reactor volume, and the details are discussed later. Preliminary runs were made to study the effect of catalyst size and the sulfur content of the feed on the product yields. The effect of catalyst size was studied for three different average catalyst diameters, viz., 3.0 f 0.5,4.0 f 0.5, and 7.0 f 0.5 111111. For each of these runs, the catalyst weight was 3 g and the temperature, 7,and d were kept fixed at 1073 K, 0.18 s, and 0.8 kg of steam/(kg of naphtha), respectively. For the range investigated, there was no effect of catalyst size on the product yields. The de-
Table 111. Composition of the Gaseous Product Mixture (wt % Feed) from the Steam Cracking of Naphtha (T = 1073 K, 0.18 s; 6 = 0.8 ka of Steam/(kg of NaDhtha)) thermal pyrolysis catalytic pyrolysis feed catalyst size, mm product yields, wt % feed carbon monoxide carbon dioxide methane ethane ethylene acetylene propane propylene n-butane isobutane 1-butene 2-butenes isobutylene + 1,3-butadiene other gases
0.2 10.2 1.7 25.3 0.4 0.4 12.3 0.4 0.3 2.3 2.0 5.2 1.6
total gas yield, wt %
63.1
T
desulfurized naphtha
desulfurized naphtha 3.0 f 0.5
desulfurized naphtha 7.0 0.5
straight-run naphtha 3.0 f 0.5
0.8
2.2 0.3 14.0 2.6 33.6 0.4 0.6 14.5 0.4 0.3 2.0
2.0 1.0 13.8 2.7 32.5 0.3 0.8 14.2
1.6
5.4 1.5
1.9 5.7 1.7
79.4
79.5
0.0 0.0 15.2 2.4 33.5 0.5 0.5 14.0 0.3 0.3 1.6 1.0 5.9 0.9 76.1
*
0.6
0.2 2.1
=
Ind. Eng. Chem. Res., Vol. 31, No. 1,1992 149 Table IV. Effect of Operating Conditions on the Composition of the Gaseous Product (wt % Feed) for the Catalytic Pyrolysis of Naphtha TC5 run TC2 TC6 sc1 sc4 DC1 DC4 temperature, K 1023 1103 1123 1073 1073 1073 1073 space time, s 0.17 0.19 0.19 0.13 0.3 0.18 0.18 0.8 steam to naphtha weight ratio, kg/kg 0.8 0.8 0.8 0.8 0.48 1.35 product yields, wt % feed carbon monoxide carbon dioxide methane ethane ethylene acetylene propane propylene n-butane isobutane 1-butene 2-butenes isobutylene + 1,3-butadiene other gases total gas yield, wt % go
c 8
0.7 0.1 8.9 2.2 25.5 0.1 0.7 13.0 1.0 0.2 5.7 5.4 5.0 6.1 74.6
4.0 0.7 17.2 2.4 35.5 0.7 0.5 13.5 0.2 0.3 1.1 1.5 5.3 0.9 83.8
6.5 2.5 20.1 2.2 39.1 1.0 0.3 12.1 0.3 0.5 1.1 4.6 90.3
1.2 0.7 12.7 1.3 31.3 0.2 0.6 14.0 0.4 0.3 3.1 2.1 4.7 1.1 73.7
5.2 3.1 16.6 0.8 34.7 0.3 9.7 14.2 0.3 0.2 0.6 0.3 6.3 0.8 84.1
1.3 0.2 12.8 2.3 31.8 0.3 0.8 14.1 0.3 0.2 2.3 1.8 5.0 1.5 74.7
5.9 2.5 15.1 1.7 34.5 0.4 0.7 14.3 0.2 0.1 0.7 0.6 6.1 0.7 83.5
.0.18?0.01s =0.8kglkg
0 with catalyst
A without cotalyst
.Ot
'0973
1023
O
1073
1123 d
Temperature, H
Figure 3. Effect of temperature on overall gas yields.
tailed product distribution using an average catalyst size of 3.0 or 7.0 mm is given in Table 111. These results show that, in this size range, the catalyst was not influenced by diffusional limitations and, for all subsequent runs, a catalyst with an average size of 3.0 f 0.5 mm was used. To check whether the sulfur content of the feed had any effect on the product yields, runs were made with desulfurized or straight-run naphtha, at otherwise identical conditions. Qpical results at a temperature of 1073 K are shown in Table 111. As can be seen from Table 111, the presence of sulfur in the feed significantly reduces the formation of carbon oxides without significantly affecting the yield of the other products. In order to have measurable differences in the yields of carbon oxides during thermal and catalytic pyrolysis for the main pyrolysis runs, the feed used was desulfurized naphtha. Effect of Temperature. The effectiveness of the catalyst was studied in the temperature range of 973-1123 K by measuring the overall gas and liquid yields together with the yields of the individual C1-C4 gases and carbon oxides. At each temperature, runs were made at identical conditions for thermal as well as catalytic pyrolysis. For these set of runs, the weight ratio of steam to naphtha, 6,
Temperature, K
Figure 4. Effect of temperature on methane, ethane, and ethylene yields.
was kept fixed at 0.8 kgjkg. Due to the axial temperature profile, a single reaction temperature for a run is not defined and the temperature mentioned for each run refers to the reference temperature (equal to the catalyst zone temperature) for which the equivalent reactor volume was evaluated (Kumar and Kunzru, 1985). Since the space time depends on the equivalent reactor volume, at different reference temperatures, 7 was not exactly the same but varied between 0.17 and 0.19 s. However, the space times for the catalytic and the corresponding thermal pyrolysis runs were identical. As expected, the major gaseous products, for both thermal and catalytic pyrolysis, were methane, ethylene, propylene, C4 olefinic gases, and 1,3-butadiene. The variation of the total gaseous product with temperature is shown in Figure 3 whereas the yields of CHI, C2H4,C&, and 1-butene are shown in Figures 4 and 5. . The detailed product distribution for some catalytic runs is given in Table IV. As can be seen from Figure 3, the gas yield increased significantlydue to catalyst addition. The effect of the catalyst diminishes at higher temperatures. For
150 Ind. Eng. Chem. Res., Vol. 31, No. 1, 1992 t = 0 18 z0.015 6 =0.8kglkg 0 CO
ACO2
y i e l d with catalyst 13
I,
18
f
s
L
Temperature, K Temperature, K
Figure 6. Effect of temperature on yields of carbon oxides.
Figure 5. Effect of temperature on propylene and 1-butene yields.
instance, at 973 K, the gas yields, which are a measure of the naphtha conversion, were approximately 22 % higher for catalytic pyrolysis, whereas, at 1123 K, the gas yields were only increased by 6% due to catalyst addition. This indicates that the activation energy for catalytic pyrolysis is lower than for thermal pyrolysis, as confirmed by the kinetic analysis later. For both catalytic and thermal pyrolysis, yields of methane and ethylene increased with temperature (Figure 4), whereas the yields of propylene and 1-butene pass through maxima with increasing temperature (Figure 5). At the same operating conditions, yields of methane, ethane, and ethylene were significantly higher during catalytic pyrolysis. The maxima in the yields of propylene, 1-butene,and 2-butenes (refer to Table N) occur at a lower temperature during catalytic pyrolysis. For instance, the yield of propylene was a maximum between 1073 and 1088 K for catalytic pyrolysis, whereas, during thermal pyrolysis, the maximum yield was obtained between 1100 and 1110 K. The concentrations of these olefinic gases decrease due to the secondary pyrolysis and polymerization reactions. For the same gas yield, the product distributions, for thermal and catalytic pyrolysis were nearly the same. For example, for the thermal pyrolysis run at 1023 K (overall gas yield, 58.2 wt %), the CH4, CzH4,C3H6,1-butene, 2butene, and 1,3-butadiene + isobutylene yields were 5.9, 16.1,9.1,6.8,6.0, and 4.3 wt 70, respectively, compared to 5.0, 16.1, 10.4,4.8, 3.9, and 4.2 wt %, respectively, for the catalytic run at 973 K (overall gas yield, 55.1 wt 70). The yields of carbon monoxide and carbon dioxide continuously increased with temperature for both catalytic and thermal pyrolysis runs (Figure 6). Addition of the catalyst significantly increased the production of CO and COP,in agreement with the trends reported for catalytic pyrolysis of n-hexane (Lemonidou and Vasalos, 1989). Kikuchi et al. (1985), for the catalytic pyrolysis of naphtha on calcium aluminate catalyst, found that the C02yield increased with temperature until 1123 K and decreased above this temperature. They attributed this decrease to the coking of the catalyst. No maxima in COz yields were observed in this study, possibly because the highest reaction temperature was 1123 K. For catalytic pyrolysis of naphtha at 1093 K, Tomita et al. (1977) found the CO
zol 0.10
0 with catalyst 0 without catalyst
0.15
0.20
0.25
0.30
0 5
space time, s
Figure 7. Effect of space time on overall gas yields.
and COPyields to be 1.9 and 13.5 w t 9%,respectively. The CI-C4yields reported by theae authors are nearly the same as those obtained in this study a t T = 1123 K, 6 = 0.8 kg of steam/(kg of naphtha), and 7 = 0.19 s (runTC6, Table IV). The significantly higher COPyield in their study is most probably due to the higher gasification activity of the catalyst prepared by them. Effect of Space Time. The effect of space time was studied in the range of 0.13-0.3 s, keeping the temperature and weight ratio of steam to naphtha fiied at 1073 K and 0.8 kg/kg, respectively. For both catalytic and thermal pyrolysis, overall gas yields increased gradually with space time, tending to level off at higher space times (Figure 7). The increase in gas yield due to catalyst addition was leas appreciable at higher space times. Depending on the stability of the hydrocarbon and the secondary reactions, yields of the various products either increased, decreased, or showed a maximum with increasing space time. Thus, methane yield increased monotonically with space time, whereas the ethylene yield tended to level off at the higher
10 T = 1073K 6 = 0.8kg/kg 0 co yleld with catalyst A C Q 0 CO A cO2
8V
8 z c
81
19
I
18
18
without n
”
19
-
P
01 0.10
I
0.15
I
I
I
0.20 0.25 Space time, s
0.30
0.35 5
Figure 8. Effect of apace time on methane and ethylene yields.
--
T
o
1073K
6 = O.Ekg/kg
0
C3Hg yield with
A
I-C& C3He 1-C&
0
A
9
n 10
catalyst
Space time, s
Figure 10. Effect of apace time on yields of carbon oxides.
90
without I
m 8,
T E 1073K I; 5 0.18~0.01s 0 with catalyst without catalyst
0 Space time, s
0.4
0.6 0.8 1.0 1.2 Weight ratio of steam to naphtha, kg / kg
1
Figure 9. Effect of apace time on propylene and I-butene yields.
Figure 11. Effect of d on overall gas yields.
space times (Figure 8). The propylene yield showed a maximum (not very pronounced), whereas the yield of l-butene decreased with increasing time (Figure 9). As shown in Table IV (runs SC1 and SC4), yields of all C4 gases, except l,&butadiene, decreased with an increase in space time. As shown in Figure 10, with an increase in space time, more of the carbon is gasified, resulting in increased yields of CO and COP. Effect of the Weight Ratio of Steam to Naphtha. The weight ratio of steam to naphtha, 6, was varied from 0.48 to 1.35 kg/kg at a temperature of 1073 K. The flow rates of water and naphtha were adjusted such that the space time was maintained at 0.17-0.18 s. As 6 was increased from 0.48 to 1.35 kg of steam/(kg of naphtha), the yield of gaseous products for thermal pyrolysis increased from 60 to 74% and from 75 to 84% for catalytic pyrolysis (Figure 11). The influence of 6 on the yields of ethylene, propylene, and methane is shown in Figure 12, whereas the detailed product distributions for two values of 6 are given in Table IV (run DCl and DC4). The yield of
ethylene increased from 31.8 to 34.5% for catalytic pyrolysis and from 22.8 to 27.6% for thermal pyrolysis. Apparently, the decrease in the partial pressure of the reacting components does not have an appreciable effect on the yields of methane and propylene. These trends are in good agreement with the data presented by &chi et al. (1985) for catalytic pyrolysis of naphtha at 1173 K. For both catalytic and thermal pyrolysis, CO and COzyields increase with increasing 6 (Figure 13). An increase in 6 enhances the formation of CO and COz, most probably due to the higher rates of steam reformation and/or coke gasification reactions. Product Selectivities. The selectivity of a product is defined as the mass (or moles) or a component formed per unit mass (or mole) of naphtha cracked. Due to the complex nature of the feedstock, it is very difficult to determine the exact amount of naphtha converted in any run (van Damme et al., 1981). It has been reported (Kumar and Kunzru, 1985; Hirato and Yoshioka, 1973) that, at low conversions (
