Artificially Metals-Poisoned Fluid Catalysts. Performance in Pilot Plant Cracking of Hydrotreated Resid E. Thomas Habib, Jr.,*' Hartley Owen, Paul W. Synder, Carl W. Streed, and Paul B. Venuto' Mobil Research and Development Corporation, Research Department, Paulsboro, New Jersey 08066
Selectivity patterns for cracking hydrotreated resid at 1000 O F and 3-5 s contact time over equilibriuni catalyst artificially poisoned with nickel and vanadium were studied. Conversion dropped only slightly, leveling out rapidly with increasing metals level, while C5+ gasoline yield showed an initial decline of 3%, but then remained constant. Coke on fresh feed increased sharply, nearly doubling at the 900 ppm level, Put showed little increase thereafter. Hydrogen production increased sharply, but tended to flatten out after the 900 ppm level. The maxima in the gasoline yield occurred at lower conversion for the metals-poisoned catalysts than for the nonpoisoned base catalyst; octane numbers (R 0) with the poisoned catalysts (70-85 vol % conversion) were 87.5-90.0, about one R 0 unit higher than the base. PONA analysis of C6+ gasoline showed no unusual trends. Total dry gas yields (wt %) showed little change due to metals poisoning, but composition changed.
+
+
Introduction As the demand for gasoline and light fuel oil increases while crude oil becomes both more expensive and less available, there is interest in processing the heavier crude fractions. Conversion of high molecular weight cuts into lighter, premium products necessarily involves a cracking step. However, the concentration in these fractions of nickel, vanadium, and nitrogen impurities, which poison cracking catalysts, complicates matters. There has been interest recently in hydrotreating heavy fractions such as atmospheric resid and heavy gas oil (Hemler and Vermillion, 1973; Hildebrand et al., 1973; McCullouch, 1975; Haunschild et al., 1975; Milstein et al., 1976; Ondish et al., 1976; Owen et al., 19761, to remove these contaminants, with subsequent catalytic cracking. However, even severe hydrotreating may not completely remove the nickel and vanadium. Any metals in the FCC feed usually deposit on the cracking catalyst, and even small concentrations will eventually build up to high catalyst metals levels. The effects of metals on amorphous cracking catalysts have been reported (Connor et al., 1957; Grane et al., 1961). Cimbalo e t al. (1972) have published an excellent comparison of the effects of metals poisoning on amorphous and zeolitic catalysts in the cracking of Mid-Continent gas oil. We report the effect of increasing catalyst metals level on the conversion/selectivity patterns in the catalytic cracking of hydrotreated atmospheric resid at 1000 O F and short contact times in a riser FCC pilot plant. Commercial equilibrium fluid zeolite cracking catalyst, artificially poisoned to several metals levels, was employed. Data on conversion, yield structure, and product quality are included, with some mechanistic interpretation. Experimental Section Charge. The feedstock was prepared by hydrogenation of a light Arabian atmospheric resid. Inspections for this hydrotreated (HDT) resid are shown in Table I. Catalyst. The metals-contaminated catalysts were all prepared from the same base commercial equilibrium catalyst by impregnation with a toluene solution of nickel and vanadium naphthenate. Three solutions were used to produce catalysts metals-poisoned to 510,870, and 2080 ppm of nickel To whom inquiries about this paper may be addressed at Mobil Research and Development Corporation, Field Research Laboratory, 3600 Duncanville Road, Dallas, Texas 75236.
equivalent, where nickel equivalent equals 100% of the active nickel plus 25% of the active vanadium, as indicated by the work of Cimbalo et al. (1972). The toluene was evaporated, and the catalysts were calcined for 6 h at 1000 O F in air. All catalysts were free of carbon before use and were not reused. Inspections of the metals-poisoned catalysts and the nonpoisoned base catalyst are shown in Table 11. Apparatus. A 30-ft riser-type FCC pilot plant was operated isothermally at 1000 OF. A flow diagram of the equipment is shown in Figure 1. Hot catalyst flowed downward from a pressurized storage hopper through an orifice into a mixing zone where the oil charge, preheated to 790 O F , was injected. The hot catalyst vaporized the oil, and the mixture swept through the riser-reactor and into a catalyst receiver. Liquid salt baths maintained the oil preheater and reactor loops a t constant temperature. Steam flowing upward in the catalyst receiver fluidized the spent catalyst and stripped it of interstitial vapors. Reaction products (except coke) passed through a filter into a conventional product collection/measurement train, and were then analyzed. Coke was determined by combustion of spent catalyst and analysis of off-gas in a modified Orsat apparatus. Operating conditions and product yields are shown in Table 111. Conversion is defined as 100 minus the cycle oil above a C5 -356 OF a t 90 vol % ASTM gasoline. Reaction temperature, oil partial pressure, and catalyst residence time were held constant a t 1000 OF ( f 2 ) , 24 ( f 5 ) psia, and 4.0 (f0.8)s, respectively. Conversion was varied by changing the catalyst/oil (wt/wt) ratio (C/O).
Results Catalyst Activity and Feed Crackability. The hydrotreated (HDT) resid was a highly crackable material, as expected from its high hydrogen content (Wollaston et al., 1971; Pierce et al., 1972; Hemler and Vermillion, 1973; Hildebrand et al., 1973; Ritter et al., 1974). The response of conversion to increased C/O with all four catalysts was about +8 vol % for doubling the C/O in the 70-80 vol % conversion range. This is similar to the results obtained by Owen et al. (1976) for riser FCC cracking of a distillate gas oil of slightly higher hydrogen content. Similarly, a H D T Kuwait stock has been observed (Hildebrand et al., 1973) to crack with distillate-like characteristics. Catalyst activity, defined as conversion a t a given C/O, was only slightly lowered by metals poisoning (see Figure 7 below). Yield Structure. Yield of all dry gas (C3-) components increased with conversion (Table 111)as Pohlenz (1963) obInd. Eng. Chem., Prod. Res. Dev., Vol. 16, No. 4, 1977
291
REOdLlTOR
CAT~LYST LOADiNG CIRCUIT
Table I. Properties of Hydrotreated Resid Feedstock
SIMPLIFIED FLOW DIAGRAM OF BENCH-SCALE R I S E R FCC PILOT PLANT
CATALYST HOPPER
LV
i -i PRO0"CT SEPARATION TRAIN
CTION: h1"S REHENT
-
;;;;& Dill"
+r::.,I.PG
Figure 1. Simplified flow diagram of bench-scale riser FCC pilot plant.
Density a t 60 O F , g/cm3 ("API) Refractive index, 70 "C Molecular weight (V.P.) Pour point, O F Conradson carbon residue, wt % Hydrogen, wt % Sulfur, w t % Total nitrogen, wt % Basic nitrogen, ppm Nickel, ppm Vanadium, ppm Iron, ppm Molecular type, wt % (mass spec.) Paraffins Naphthenes Aromatics Distillation, A.S.T.M., "F IBP 10, vol % 30, vol % 50, vol % 70, vol % 90, vol %
0.9065 (24.6) 1.4953 524 60 2.6 12.6 0.27 0.076 179 0.7 1.9 0.7
23.8 26.2 50.0 639 724 771 848 937 1062
Table 11. Inspections of Catalysts Catalyst number
I
80
I
I
90
I
J 100
CONVERSON, VOL %
Figure 2. Molar expansion against conversion, vol%, for riser cracking of HDT resid over metals poisoned catalysts. 0, 0 ,A , X = 80, 510, 870, and 2080 ppm active metals on catalyst, Ni equivalent. served. Total dry gas (wt %) did not change significantly as catalyst metals level increased, but composition shifted greatly. Moderate increases in dry gas with increasing metals level were observed by Cimbalo et .al. (1972). In our study, as catalyst metals level increased, yield of hydrogen increased sharply, methane increased slightly, and propylene decreased. The other dry gas components showed no change. Similarly, Cimbalo et al. (1972) noted sharp increases in the H&H4 ratio with increased metals poisoning. Although this resulted in little change in total dry gas yields on a weight basis, it represents a substantial increase on a molar basis. In Figure 2, the total molar expansion vs. conversion with the extrapolated 100% conversion values are shown. The increase in molar expansion with increasing metals level is due almost entirely to the increased yield of molecular hydrogen. As expected (Pohlenz, 1963), total Cq yield increased sharply with conversion for all catalysts, increasing nearly 5 vol % for a 10 vol % increment in conversion. This increase was largely due to butane (both iso- and n-)since neither the yield of butene nor the ratio of isobutane to n-butane changed greatly with conversion. The effect of increasing catalyst metals level had a different effect. Total C4 yield decreased with increasing metals level (at a given conversion). A similar trend, but not as pronounced, was observed by Cimbalo et al. (1972). This decrease was primarily in butene; total butane dropped only slightly, but the isobutane to normal butane ratio decreased substantially. 292
Ind. Eng. Chem., Prod. Res. Dev., Vol. 16, No. 4, 1977
Metal, ppm Nickel Vanadium Copper Iron Active metal, ppmb Nickel Vanadium Total, Ni equiv., ppmC Fluid activity indexd
Metal poisoned 2 3
Basea
1
190 240 10 1200
530 580
400 80 420 80 510 67.6 64.6 60
790 1000
1300 3800
660 1170 840 3640 870 2080 60.6 53.8
a Equilibrium commercial cracking catalyst containing 15%wt ReY zeolite and burned carbon free. S.A. 172 m2/g;packed density, 0.74 g/cm3; pore volume 0.64 cm3/g. Calculated as 33.3% of the original Ni and V plus 100% of the artificially added Ni and V. Calculated as 100% of the active Ni plus 25% of the active V, in accordance with recent findings of Cimbalo, et al. (1972). Defined as the conversion to gasoline (356 "F at 90 vol % A.S.T.M. distillation) and lighter when cracking a Light East Texas Gas Oil at 2C/O, 850 O F , 6 W.H.S.V. for 5 min on-stream time in a fixed fluidized bed reactor.
Total C5 yield increased with conversion, but less sharply than did the total C4 yield. This was so because although the yield of pentane increased, the yield of pentene decreased. Increasing catalyst metals level had no effect on the total yields of pentene or pentane. However, the ratio of isopentane to n-pentane, which was little affected by change in conversion, was substantially lower with the metals-poisoned catalysts than with the unpoisoned base. Gasoline yields increased, generally reached a maximum, and then decreased as conversion increased (Figure 3). Gasoline yields a t a given conversion were lower with the metals-poisoned catalysts than with the unpoisoned, and the maximum in the gasoline yield profile shifted t o a lower conversion level for the metals-poisoned catalysts. A similar pattern with metals poisoning was observed by Cimbalo et al. (1972) for the cracking of Mid-Continent stock. However, there did not appear to be any difference between the three poisoned catalysts in our study. In Figure 3, one curve is drawn for the unpoisoned catalyst and another for the most severely poisoned catalyst; the data for all poisoned catalysts appear
Table 111. Operating Conditions and Yield Structure for Riser Fluid Catalytic Cracking of HDT Resid Catalyst metal level, ppm, Ni equiv 80Conditions Reactor temperature, OF Catalyst inlet temperature, O F Catalyst/oil ratio, wt/wt Catalyst residence time, s Oil partial pressure, inlet, psia Carbon on spent catalyst, wt % Yields (no loss basis)a Conversion, vol % b Hydrogen, wt % Methane, wt % Ethylene, wt % Ethane, w t %
510-
-870-
-2080-
~1312 1187 1104 1312 1187 1156 1104 1312 1156 1104 1312 1156 1104 3.0 5.3 9.5 2.9 5.4 5.9 9.5 3.0 5.9 9.9 3.1 5.7 9.4 4.8 4.4 3.9 4.4 3.9 3.9 3.3 4.3 3.8 3.3 4.1 3.6 3.2 27.7 24.8 19.2 28.1 26.4 25.4 19.2 28.9 24.8 19.8 28.9 24.9 19.8 1.166 0.983 0.635 1.654 1.388 1.298 1.124 1.897 1.413 1.072 2.080 1.447 1.193 72.35 81.37 84.77 71.06 79.94 80.90 85.31 70.26 79.90 83.96 72.44 79.87 84.65 0.06 0.05 0.05 0.34 0.45 0.48 0.57 0.50 0.58 0.67 0.61 0.69 0.79 0.80 0.92 1.04 0.74 0.90 0.91 1.02 0.78 0.94 1.06 0.86 0.98 1.16 0.61 0.66 0.88 0.61 0.70 0.71 0.83 0.67 0.73 0.83 0.62 0.65 0.75 0.65 0.66 0.68 0.58 0.60 0.65 0.64 0.61 0.62 0.65 0.64 0.65 0.71
Propylene, vol % Propane, vol %
6.88 1.76
Total C ~ ' Svol , % Butenes, vol % Isobutane, vol % n-Butane, vol%
7.93 2.63
9.21 3.83
8.64 3.43
6.34 1.71
8.64 10.56 13.04
7.78 10.23 10.88 12.07
8.50 5.27 1.21
6.53 4.56 1.28
6.73 8.38 11.00 1.85 2.92
8.81
6.23 1.55
7.82 2.41 7.43 7.17 1.92
8.24 2.63 7.22 7.66 2.25
6.97 9.62 2.65
7.65 2.35
8.07 3.24
5.51 1.61
6.76 2.47
8.05
9.99 11.31
7.13
9.23 10.70
5.92 4.49 1.35
6.05 6.36 2.05
6.81 3.93 0.99
6.62 6.46 2.07
6.83 8.87 2.48
7.32 3.38 6.42 8.12 2.54
14.99 19.03 20.65 12.37 16.53 17.13 19.24 11.75 14.45 18.17 11.74 15.14 17.08
Total C ~ ' Svol , %
6.23 5.06 0.98
Pentenes, vol % Isopentane, vol % n-Pentane, vol %
4.59 7.64 1.51
3.50 9.39 2.11
5.74 4.81 1.17
5.01 6.69 1.78
4.99 7.18 2.04
3.95 8.56 1.98
5.57 4.99 1.26
4.46 6.15 1.74
4.15 8.21
2.05
5.49 4.11 0.73
5.00 6.55 1.83
4.25 7.69 1.99
12.27 13.74 15.01 11.71 13.47 14.21 14.49 11.82 12.35 14.42 10.33 13.38 13.92
Total Cg's, vol % CS+Gasoline, ~ 0 1 % ~ Cycle oil, vol % Coke, wt % Gasoline efficiency, vol % Gasoline octane number, R 0'
+
48.52 27.65 3.70 84.0 86.7
50.32 49.37 48.12 48.06 18.63 15.23 28.94 20.06 5.63 6.67 5.19 8.10 78.7 76.0 84.2 77.0 86.2 87.7 87.3
47.08 19.10 8.28 75.8 88.0
45.59 14.69 11.75 70.4 90.0
45.30 29.74 6.04 81.3 87.6
49.90 20.10 9.04 77.9 88.4
45.27 16.04 11.71 71.1 89.7
48.76 27.56 6.94 81.6 88.2
48.28 20.13 8.98 77.2 87.9
46.25 15.35 12.39 71.1 90.0
a Recoveries on feed averaged 93 w t %. Based on Cs to 356 O F at 90 vol % gasoline. c Clear research octane number on Cg to 356 OF at 90 vol % gasoline.
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ap 8 -
d
76-
Y
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tA
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ao
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a5
CONVERSION, VOL %
5 4-
31
'
70
I
75
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80
I
85
I
CONVERSION, VOL%
Figure 3. Cg+ gasoline yield, vol %, against conversion, vol %, for riser cracking of HDT resid over metals poisoned catalysts. 0,0 ,A, X = 80, 510, 870, and 2080 ppm active metals on catalyst, Ni equivalent.
Figure 4. Coke yield, wt %, against conversion,vol %, for riser cracking of HDT resid over metals poisoned catalysts. 0, 0 , A , X = 80, 510, 870, and 2080 ppm active metals on catalyst, Ni equivalent.
to fit the curve for the most severely poisoned catalyst. Gasoline efficiency (defined as the ratio of the vol % gasoline yield to the vol % conversion) decreased as conversion increased, and a t a given conversion it was lower for the poisoned catalysts than for the base. As shown in Figure 4, the characteristic (Pohlenz, 1963) increase in coke yield with extent of conversion was observed.
Coke yield (at constant conversion) increased sharply as equivalent nickel on catalyst increased from the base to the first level (510 ppm), showed some further increase as the metals level increased to 870 ppm, but did not increase further when the metals level was raised to 2080 ppm. Coke on catalyst (Table 111) increased sharply with catalyst metals level, but decreased as conversion (and C/O)increased. Ind. Eng. Chem., Prod. Res. Dev., Vol. 16, No. 4, 1977
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, _-
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OLEFINS VOL % I
1
ARMIATICS VOL %
25
-
COMBINED HYDROGEN WT %
O '
I
I
I
75
80
85
1
paraffins, olefins, naphthenes, and aromatics, and wt % combined hydrogen against conversion, ~ 0 1 %0, . 0 ,A , X = 80,510,870, and 2080 ppm active metals on catalyst, Ni equivalent.
+
Product Quality. Clear research octage numbers (R 0) for the C5+ gasolines were in the range of 86 to 90, tending to be about 11/2 numbers higher for the metals-poisoned catalyst than the base, possibly reflecting in part slightly higher olefin and slightly lower paraffin content (Figure 5). Octane numbers tended to be about 2 numbers higher when conversion was above that of maximum gasoline yield (Figure 3), possibly the result of increased aromaticity (Figure 5). Densities of the C5+ gasolines varied only slightly with conversion or catalyst metals level, ranging from 0.742 to 0.756, and averaging 0.747 g/cm3 (60 OF). Molecular weight of c6+gasolines averaged 108 f 4 and were independent of catalyst metals level but tended to decrease with increasing conversion. However, the total C5's also increased with conversion and this compensated, making the C5+ gasoline molecular weight independent of conversion level. Molecular compositions for the c6+ gasolines are plotted vs. conversion in Figure 5. As conversion increased, the following changes in molecular composition took place: paraffins increased initially but leveled out as conversion approached 85 vol %; olefins dropped sharply; naphthenes decreased slightly; aromatics increased substantially; and hydrogen content decreased. Similar trends in gasoline P, 0,N, A with conversion have been reported by Owen et al. (1976) for riser cracking of HDT heavy vacuum gas oil. Comparing the unpoisoned catalyst to the most severely poisoned catalyst, we note the following gasoline compositional differences at the higher catalyst metals level: paraffins and naphthenes were slightly lower, while olefin content was higher; there was no difference in aromatics or hydrogen content. Gasoline compositional changes due to increased metals on catalyst were much smaller than those due to increased conversion. c6+ gasoline vol % olefins, for example, increased by only 5% (at the 70-75% conversion level) due to catalyst metals poisoning, but decreased by 15% for a 15 vol % increase in conversion. Molecular compositions of total cycle oils are plotted vs. conversion in Figure 6. Lines have been drawn through data Ind. Eng. Chem., Prod. Res. Dev., Vol. 16, No. 4, 1977
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HYDROGEN WT %
e
P I
GRAVITY
1.1 I5
5
' s o - 0 ,
I
70
75 80 CONVERSION. VOL%
85
Figure 6. Molecular composition of total cycle oil, wt % paraffins,
C O N V E R W . VOL%
Figure 5 . Molecular composition of depentanized gasolines, vol %
294
'p'
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naphthenes, aromatics, and combined hydrogen, and A.P.I. gravity against conversion, ~ 0 1 %0, . 0 ,A , X = 80, 510,870, and 2080 ppm active metals on catalyst, Ni equivalent.
points for the unpoisoned catalyst and for the most severely poisoned case. All cycle oils consisted primarily of aromatics, which increased from 70 to 90 wt % as conversion increased from 70 to 85 ~ 0 1 %Paraffins . and naphthenes both decreased from 15 to 5 wt % for this conversion change. Similar trends were observed in cracking of HDT heavy vacuum gas oil (Owen et al., 1976). This sharp increase in aromaticity is consistent with lower hydrogen contents and A.P.I. gravities, which decreased from 10 to 8%,and from 15 to 5O, respectively. Increasing catalyst metals level had some effect on cycle oil composition: paraffins remained constant, naphthenes decreased, and aromatics increased, but the changes were too small to measurably affect the hydrogen content or the A.P.I. gravity, both of which were independent of catalyst metals level.
Discussion Yield Structure vs. Conversion. Trends in light ends proportions as conversion increased from 72 to 85% are shown in Table IV. C r C 4 olefins were generally found in greater than equilibrium concentration, reflecting the fundamental cracking mechanism of carbonium ion formation followed by 0-scission to form an olefin (and a new carbonium ion), as shown by Thomas and McNelis (1967). c 2 - C ~olefins all decreased with increasing conversion, the result of (secondary) hydrogen transfer processes well known in catalytic cracking (Voge, 1958; Shephard et al., 1962; Nace, 1969; Eastwood et al., 1971). Initial high isobutane levels so characteristic of zeolitic cracking (Nace, 1969) showed only slight drain-off (by isomerization) with increasing conversion, remaining at over twice equilibrium values even at the 85% level. The trends shown in Table IV are in striking agreement with those observed by Owen et al. (1976) for riser cracking of a hydrotreated heavy vacuum gas oil under nearly identical conditions. The trends in gasoline P, 0, N, A analysis in Figure 5 are consistent with patterns of intermolecular hydrogen transfer in the presence of zeolites, where, as Nace (1969) has proposed, khydrogen transfer/k~.scission is far greater with zeolites than with amorphous silica-alumina type. Eastwood et al. (1971) point
Table IV. Changes in Proportions of CZ-C~ Fractions with Increasing Conversion over Base Catalyst Equilibrium, Variable
Range
96 CzH4 in total Cp
48 80 57 50 82
% C3H6 in total Ca % C4Hs in total Cq % CsH10 in total Cg % i-Cq in butanes
+
+
57 71 33 23 80
1000 O F a
8 27 456 33‘
out how such hydrogen redistribution occurs with great facility in feedstocks where hydrogen has been catalytically introduced. Such reactions include 2 olefin aromatic t paraffin, olefin naphthene aromatic paraffin, etc; note that polymer and coke may also be formed (i.e., on the catalyst) in addition to monomeric aromatics. Total cycle oil, by definition, decreases as conversion increases. Weekman (1968) observes that in the cracking of such a complex mixture as gas oil, different classes of molecules crack a t different rates, and hence, reactivity itself is a function of conversion level. We see this in Figure 6, where the more reactive, hydrogen-rich paraffins and naphthenes are cracked away, leaving a concentration of less crackable, more chemically resistant aromatics. The trend toward higher density and greater hydrogen deficiency naturally follows. Yield Structure vs. Metals Level. The effect of metals poisoning on catalyst activity and selectivity is summarized in Figure 7. The decline in catalyst activity (corresponding to the drop in conversion a t constant C/O) (Figure 7A) as metals level increases is somewhat less than reported by Cimbalo et al. (1972) for the cracking of a Mid-Continent gas oil. However! that work was conducted at much longer catalyst residence times, and possibly the lower conversions obtained with the metals-contaminated catalysts at least in part reflect the time-dependent catalyst deactivation by coke buildup (Weekman, 1969). To relate these laboratory results directly to commercial performance, adjustments must be made for differences in average age of catalyst metals, residual carbon level, stripper efficiency, etc. However, even without adjustment, the directional trends found in this study can be used to assess some consequences of metals poisoning in commercial units. The increased hydrogen yield observed with metals poisoning would have considerable effect on the performance of the wet gas compressors (Connor et al., 1957; Grane et al., 1961; Pohlenz, 1963; Cimbalo et al., 1972). These compressors are usually centrifugal and the high hydrogen yields cause problems not only because of the large volumes, but also because the low gas density limits the maximum pressure ratio attainable per stage. Similarly, coke yield is of utmost importance in commercial FCC units since they operate in a heat-balanced mode (Connor et al., 1957; Grane et al., 1961; Pohlenz, 1963; Cimbalo et al., 1972) and it is necessary for coke yield to be within a rather narrow range to fulfill the constraints of the energy balance. Cimbalo et al. (1972) classify FCC coke yields into four categories: (1)“catalytic” coke, which arises in association with acid-catalyzed cracking reactions and is residence time-dependent as shown by Voorhies (1945);(2) “contaminant” coke, which derives from catalytic action of metals poisons such as nickel; (3) “Conradson” coke, which is assumed to correlate directly with Conradson carbon analysis, and (4) “Cat-to-oil” coke, which in essence measures unstripped but potentially strippable hydrocarbons, which are not really coke, but because of nonideal stripping in the commercial process, are also burned in the regenerator.
-
+
75
, 1
PRODUCT SELECTIVITY
YIELDS AT 80 VOL% CONVERSION
60
Kirby (1955). * Mixed n-C4’s. Lawrence and Rawlings (1967).
+
AT 5 ‘00
C; GASOLINE VOL 90
-
HWpT GAS X
0
pyf
0 500 1000 15W 2000 ACTIVE METAL ON CATALYST, PFW NI EWIVALENT
Figure 7. Catalyst activity, vol% conversion at 5C/O, and selectivity, yield of Cs+ gasoline, vol %, coke, wt %, and H2 Gas, w t %, against active metals on catalyst, ppm Ni equivalent. Because of highly efficient stripping in our pilot unit, the unstripped hydrocarbons are negligible. Catalytic coke, as a percent of feed, is usually directly proportional to catalyst/oil ratio (other conditions being constant). We would prefer to use the term “additive carbon” (Cadd) for the third type of coke in the Arco classification. By this we mean that fraction of the coke which is independent of catalyst/oil ratio or catalyst type, instead correlating with the basic nitrogen and average molecular weight of the feedstock, as well as the Conradson Carbon Residue (Cimbalo et al., 1972;Masologites and Beckberger, 1973). By plotting coke on fresh feed vs. catalyst/oil ratio, the intercept by extrapolation back to zero catalyst/oil may be interpreted as CAdd. Using this technique (with minor adjustments for differences in oil partial pressure and residence time), a value of about 2 wt % CAdd was determined for the HDT resid. To compare the yield of “contaminant” coke due to increasing metals level for the high molecular weight HDT resid (with significant “additive” carbon) to that obtained with a “clean” Mid-Continent gas oil, CAdd must be subtracted. When this is done, we find that the increase in active metals level from the base to the 870 ppm level doubles the “catalytic contaminant coke” yield. This is in reasonably good agreement with the findings of Cimbalo et al. (1972) over a similar conversion range (note that our nickel equivalent = their vanadium equivalent/4). As catalyst metals level increased to 2000 ppm, yields of propylene, butene and isobutane decreased by 2,3, and 2 vol %, respectively, at constant conversion. Since these C3 and Cq hydrocarbons represent a significant potential increment of gasoline by alkylation (McGovern, 1972), and particularly for butene, since butene alkylate quality and yield are superior to those from propylene (Anderson, 1974), these yield declines may be considered as gasoline yield loss. Thus, for the 80% conversion level, we calculate (adding the appropriate amount of outside isobutane), a 33% loss in alkylate in going from the base to the most severely poisoned catalyst. Our data show that the hydrogen and coke yield increases are nearly parallel (Figure 7C,D). A similar observation was made by Masologites and Beckberger (1973) for FCC cracking of a high-sulfur U.S. crude over a metals-contaminated amorphous catalyst. Cimbalo et al. (1972) found that increased yields of dry gas and coke counterbalanced gasoline yield loss.
+
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In 3ur study, a large increase in coke and small increases in hydrogen and light gas appear to counterbalance decreases in propylene, butene, and isobutane as well as a modest decrease in gasoline. It is well known for both amorphous and zeolitic cracking catalysts (Connor et al., 1975; Cimbalo et al., 1972; Masologites and Beckberger, 1973) that the (effective) metals deposited on the surface superimpose their dehydrogenation activity upon cracking reactions, thereby converting to gas and carbonaceous residue some feedstock that would more profitably appear in the gasoline fraction. The shift in peak C,+ gasoline yield with increasing metals contamination (Figure 4) (also observed by Cimbalo et al., 1972, with Mid-Continent feed) suggests some gasoline recracking, although it does not necessarily rule out "non-selective" cracking of some heavier feed molecules. In particular, nickel is known to catalyze dehydrogenation and condensation reactions-including formation of carbon and hydrogen-at high temperature (Germain, 1969). Even under hydrogenolytic conditions, Taylor et al. (1965) postulated that ethane formed an unsaturated surface radical on nickel (Nikieselguhr catalyst) via a preliminary endothermic dehydrogenation-attesting to nickel's dehydrogenative power. Aromatics have been formed from C6 and C7 naphthenes (Germain, 1969), and small amounts of olefins and aromatics have been reported from paraffin reaction over N i x zeolite (Galich, 1964). Adsorbed carbon residue was also reported by McKee (1962) in a study of ethylene decomposition on clean, unsupported nickel powder at low temperatures (392 O F ) . Nizkel-catalyzed reactions such as these probably account for most of the "contaminant" coke and hydrogen formation observed in our study. Catalytic cracking generally proceeds through a carbonium ion mechanism (Voge, 1958; Oblad, 1972),and since primary and, particularly, methyl carbonium ions are the most difficult to form energetically, C1 and C2 hydrocarbons-above and beyond any minor contribution from thermal cracking (Thomas and McNelis, 1967)-are not generally produced. Thus, we attribute the small increase in methane with increasing catalyst metals level to radical-type nickel-catalyzed a-cracking that is well known in the literature (Haensel and Ipatieff, 1946,1947; Myers and Munns, 1958; Koechloefl and Bazant, 1968; Matsumoto et al., 1970; McLaughlin and Pope, 1972). Light C2-C4 olefins and paraffins, as well as gasoline range paraffins, olefins, naphthenes, and aromatics, can undergo dehydrogenation reactions catalyzed by nickel, but it is not easy to distinguish the contribution from secondary reactions vs. that from primary gas oil reactions in the present work.
Conclusions (1)At constant conversion, catalyst metals poisoning resulted in a large increase in coke, and small increases in hydrogen and light gas. These increases were counterbalanced by decreases in propylene, butene, isobutane, as well as a modest decrease in gasoline. (2) Although the increase in hydrogen was small on a weight basis, it represents a large increase on a molar basis, and would
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have serious consequences on commercial wet gas compressors. (3) Gasoline octane numbers (R 0)were approximately 11/2 units higher with the metals-poisoned catalysts, possibly due to slightly higher olefin and slightly lower paraffin content. (4) Most of the effects of metals poisoning had occurred a t an active metals level of 500 ppm Ni equivalent. Further increases in metals level had smaller incremental effects.
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Acknowledgments We thank Messrs. F. E. Harvey and A. M. Lampkin for their excellent technical assistance, Mr. E. J. Demmel for very helpful technical discussions, Dr. R. H. Fischer for preparing the feedstock, and Mr. S. M. Oleck for preparing the catalyst. Literature Cited Anderson, R. F., Oil Gas J., 72 (6), 78 (1974). Cimbalo, R. N., Foster, R. L., Wachtel, S. J., Oil Gas J., 70 (20), 112 (1972). Connor, J. E., Jr., Rothrock. J. J., Birkheimer, E. R., Leum, L. N., Id.Eflg. Chem., 49,276 (1957). Eastwood, S. C., Plank, C. J., Weisz, P. B., Proc. Eighth WorldPet. Congr., 4, 245 (1971). _ . Galich, P. N., Gutyrya, V. S., Neimark, I. E., Egorov, Yu. P., ll'in, V. G., Galubchenko, I.T., Frolova, V. W., Neftekhim., Akad. Nauk. Ukr, SSR, lnst. Khim. Vysokomolekul. Soedin., 13 (1964). Germain. J. E.. "Catalvtic Conversion of Hydrocarbons," Academic Press, London, 1969. Grane, H. R., Connor. J. E., Masologites, G. P., Pet. Refiner, 40, 168 (1961). Haensel, V., Ipatieff, V. N., J. Am. Chem. SOC., 68, 345 (1946). Haensel, V., Ipatieff, V. N., hd. Eng. Chem., 39,853 (1947). Haunschild, V. M., Chessmore, D. O., Spars, B. G., paper presented at A.1.Ch.E. Meetina in Houston. Texas. March 1975. Hemler, 6. L.,Vermillion, W. 'L.,Oil Gas J., 71,(45), 88 (1973). Hildebrand, R. E., Huling, G. P., Ondish, G. F., Oil Gas J., 71 (50), 112 (1973). Kirby, K. K., in "Catalysis," P. H. Emmett, Ed., Vol. Ill, Chapter 10, p 455, Reinhold, New York, N.Y., 1955. Koechloefl, K., Bazant, V., J. Catal,, IO, 140 (1968). Lawrence, P. A., Rawlings, A. A,, Proc. Seventh WorldPet. Congr., 113, 101 (1967). McCullouch, D. C., Oil Gas J., 73,(29), 53 (1975). McGovern, L. J., paper presented at the National Meeting of the American Chemical Society in New York, N.Y., Aug 1972. McKee, D. W.. J. Am. Chem. SOC., 84, 1109 (1962). McLaughlin, J. R., Pope, C. G., J. Catal., 26,370 (1972). Masologites, G. P., Beckberger, L. H., Oil Gas J., 71 (47), 49 (1973). Matsumoto, H., Saito, Y., Yoneda. Y., J. Catal., 19, 101 (1970). Milstein, D., Graven, R. G., Streed, C. W., Fuselier, P. C., Oil Gas J., 74 (29), 138 (1976). Myers, C. G., Munns, G. W., Jr., lnd. Eng. Chem., 50, 1727 (1958). Nace, D. M., lnd. Eng. Chem., Prod. Res. Dev., 8, 24 (1969). Oblad. A. G., Oil Gas J., 70 (13), 84 (1972). Ondish, G. F.. Frayer, J. A., McKinney, J. D., Hydrocarbon Process., 55 (7), 105 (1976). Owen, H.. Snyder, P. W., Venuto, P. B., paper presented at the Sixth International Congress on Catalysis in London, July 1976. Pierce, W. L., Souther, R. P., Kaufman, T. G., Ryan, D. F., Hydrocarbon Process., 51 (5), 92 (1972). Pohlenz, J. B., OilGas J., 61 (13), 124(1963). Ritter, R. E., Blazek, J. J., Wallace, D. N., Oil Gas J., 72 (41), 99 (1974). Shephard, F. E., Rooney, J. J., Kemball, C., J. Catal., I, 379 (1962). Taylor, W. F., Sinfeld. J. H., Yates, D. J. C., J. Phys. Chem., 6% 3857 (1965). Thomas, C. L., McNelis, E. J.. Proc. Seventh World Pet. Congr., l B , 161 (1967). Voge, H. H., in "Catalysis," P. H. Emmett, Ed., Vol. VI, Chapter 5, p 407, Reinhold, New York, N.Y., 1958. Voorhies. A,, Jr., lnd. Eng. Chem., 37,318 (1945). Weekman, V. W., Jr., lnd. Eng. Chem., Process Des. Dev., 7,90 (1968). Weekman, V. W., Jr., lnd. Eng. Chem., Process Des. Dev., 8, 385 (1969). Wollaston, E. G., Forsythe, W. L., Vasalos, J. A., Proc., Am. Pet. lnst., Div. Refining, 51, 12 (1971). I
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Received f o r review May 13, 1977 Accepted September 1,1977
