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Ind. Eng. Chem. Res. 2003, 42, 426-433
Laboratory Evaluation of Cracking Catalysts in a Fluid Bed: Effects of Bed Dynamics and Catalyst Deactivation C. P. Kelkar, Mingting Xu, and Rostam J. Madon* Engelhard Corporation, 101 Wood Avenue, Iselin, New Jersey 08830
Testing of fluid cracking catalysts in laboratory microactivity units is carried out in either fixedor fluid-bed reactors. The latter are becoming increasingly popular because of the automation they provide. This paper focuses on understanding the effects of testing protocols on product selectivities in such reactors and explains the cause of artificial conversion and selectivity effects reported previously when a variable time-on-stream protocol was used. We show that coke deposited at the onset of a run lowers the coke selectivity of the subsequent feed to the reactor. This, in turn, affects all other product yields. The appropriate method to change conversion is to vary the mass of catalyst and thus change space time. This is best practiced without introducing other artifacts by using inerts to dilute the catalyst so that the total solid content in the reactor is constant. Finally, we have shown that a small fluid bed is significantly backmixed. Therefore, as is normal for a backmixed system, when compared to a plug-flow fixed bed at constant time on stream, the fluid bed showed lower conversion at the same space time, increased yields of final products, and lowered yields of intermediate products. However, yields from a backmixed fluid-bed reactor at longer times on stream approached yields from a plugflow fixed-bed reactor operated at a shorter time on stream. 1. Introduction Accurate evaluation of fluid catalytic cracking (FCC) catalysts is important because small changes in activity or yields can impact a refinery’s operation and profits. Circulating pilot units1 provide the best small-scale simulation of refinery risers. However, these units are expensive to operate, require a large catalyst charge, and are therefore not useful for routine catalyst research. Smaller laboratory-scale testing continues to play an important role in the evaluation of cracking catalysts. The standard microactivity test (MAT), as described in ASTM 3907-87 and carried out in a fixed bed at 755 K, uses 4 g of catalyst with a 72 s oil delivery time.2 Most laboratories have modified this test2-4 to better reflect continuing changes in commercial FCC operation such as the use of higher activity catalysts, shorter contact times, and higher reaction temperatures. Currently, most laboratories operate fixed-bed units at temperatures around 800 K with oil delivery times of 30 s or shorter with modifications to improve isothermal behavior. Another type of MAT employs a fluid-bed reactor.5,6 The operational details have been described in detail by Kayser.5 Conversion in both types of reactors may be changed by varying the weight of catalyst and that of oil, with the term catalyst-to-oil ratio often being used as a determining parameter. This ratio correctly relates to conversion changes in FCC riser reactors and helps to define the cracking operation. In laboratory reactors, it does not have the proper implication as it does in the riser. And although initially it will be necessary for us to use the catalyst-to-oil ratio term to make some points, we will later revert to the rigorous and correct concept * To whom correspondence should be addressed. E-mail:
[email protected].
of space time (τ). In circulating pilot-plant or commercial riser reactors, conversion is commonly changed by changing the catalyst quantity. In a laboratory unit, either the catalyst weight or the oil weight may be changed. A third parameter specific only to laboratory units is time on stream (ts); it is the time over which oil is delivered to the catalyst and should not be confused with contact time or with τ, where
τ ) weight of catalyst/oil injection rate
(1)
Two approaches are used in laboratory testing to vary conversions. In one case, ts is kept constant and τ is varied, and in the second case, ts is varied and τ is kept constant. In either case, one may vary either the catalyst or oil weight, always ensuring that bed dynamics are constant without introducing artifacts. For example, in the first approach at constant ts, the weight of oil may be varied by changing the oil injection rate. However, this could lead to external mass transfer limitations at higher conversions. Varying catalyst weight while keeping oil injection rate constant may introduce an artifact because the bed length-to-diameter ratio of a fixed-bed reactor would change. In fact, this problem would be more pronounced in a fluid bed. One can circumvent these problems by using the catalyst dilution approach,7 where the total weight of solids charged to the reactor is constant and only the active catalyst fraction is changed. In the second approach at varying ts, the oil injection rate is constant, and one obtains different conversions by changing the weight of oil fed to the reactor while keeping the weight of catalyst constant; therefore, here τ is constant. Recently, Wallenstien et al.8 compared the performance of catalysts, with different rare earth levels, in laboratory fixed- and fluid-bed reactors to that in a circulating riser. Conversion in the fixed-bed reactor was changed by changing τ and in the fluid-bed reactor by
10.1021/ie0206251 CCC: $25.00 © 2003 American Chemical Society Published on Web 01/07/2003
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changing ts. At a constant catalyst-to-oil ratio for the 1.1% rare earth oxide (REO) and the 2.9% REO catalysts, conversion in the circulating unit was 68 and 76%, respectively, while in the fixed-bed reactor, it was 66 and 78%, respectively. In contrast, conversion in the fluid bed was almost invariant at 74 and 77%, respectively. Increasing the rare earth exchanged on the catalyst resulted in the expected higher gasoline and lower liquid petroleum gas (LPG) yield when using the fixed-bed reactor and the circulating riser. The reverse was observed for the fluid-bed reactor. It is not clear from their work whether these effects are due to changes in reactor dynamics or the testing protocol being used. Our paper discusses different approaches in testing FCC catalysts in a small fluid-bed reactor. By understanding testing protocols and their implications, we will recommend a preferred testing approach. Although this paper focuses on fluid-bed reactors, we will also use results from a fixed-bed reactor to highlight differences between the two. Finally, we will show how catalyst deactivation in a fluid-bed reactor plays a critical role in determining product yields. 2. Experimental Section 2.1. Catalyst and Feedstock. We used a commercial 1.2 wt % REO USY based FCC catalyst containing approximately 40 wt % zeolite with 0.3 wt % Na2O. The catalyst, deactivated by steam in a fluidized bed at 1060 K and a steam pressure of 101 kPa for 4 h, had a zeolite surface area, defined as surface area of pores smaller than 2 nm, of 199 m2/g, a matrix surface area of 94 m2/ g, and a unit cell size of 2.429 nm. Inert microspheres, prepared by calcining spray-dried kaolin microspheres, had an apparent bulk density and particle size distribution very similar to those of the commercial FCC catalyst. These inert microspheres, steamed at 1090 K and a steam pressure of 101 kPa for 5 h, had a Brunauer-Emmett-Teller surface area of 9 m2/g with no surface area in pores smaller than 2 nm. We used a heavy aromatic gas oil with a low Conradson carbon (CCR) of 0.22 wt % and an API gravity of 19.2 as our feed. 2.2. Fluid-Bed Reactor. We used a commercially available fluid-bed reactor manufactured by Xytel Corp. under the trade name ACE.5 The reactor consisted of a 1.6 cm i.d. stainless steel tube with a tapered conical bottom. Nitrogen, flowing from the bottom, fluidized the catalyst and also served as the stripping gas at the end of a cracking cycle. Preheated oil was fed from the top via an injector tube with its tip close to the bottom of the fluid bed. Cold flow simulations of the fluidized bed in a glass reactor of identical proportions showed that such injection conditions ensured uniform and complete contact of the oil with the entire bed. Liquid product, collected via a side arm into a syncrude receiver maintained at 260 K, was analyzed by ASTM method D 2887 using 494 and 616 K as cut points for gasoline and light cycle oil (LCO), respectively. We analyzed all gaseous products by standard gas chromatographic techniques. We treated the deactivated and stripped catalyst in air at 955 K, quantitatively measured CO2 via an IR cell, and converted the percentage of measured carbon to coke assuming that catalytic coke had a carbon-to-hydrogen ratio similar to that of benzene. We discarded the regenerated catalyst and used a fresh catalyst charge for each experiment.
Mass balances for each run were between 98 and 101%. We report yields as a weight percentage of oil fed, and define conversion as the percent yield of products boiling below 494 K plus coke yield. We obtain yields at a specific conversion by using standard regressions. 2.3. Protocols. For the variable ts protocol, we used 9 g of catalyst charge and a 1.2 g/min oil injection rate with ts varied from 56 to 150 s. The oil injection rate and weight of catalyst were chosen because these are the parameters most commonly used in recent publications on the testing of cracking catalysts in ACE units.5,6,8 We also evaluated this protocol with other physical parameters, i.e., 12 g of total solids and a 2.8 g/min oil injection rate, and arrived at similar conclusions. For the constant ts protocol, we used a 2.8 g/min oil injection rate and 12 g of total solids. To understand the effects of deactivation on product yields, we carried out experiments at several different ts ranging from 15 to 90 s. We varied oil injection rate from 1.7 to 2.8 g/min while holding space time and ts constant. No change in conversion was observed at the higher oil injection rates. We used a 2.8 g/min oil injection rate to provide a high linear velocity through the bed to minimize the effects of external mass transfer limitations, while the choice of 12 g of total solids ensured the most uniform gas oil contacting and allowed us to run low activity catalysts at reasonable conversions. As described in the previous section, we changed conversion by changing the fraction of active catalyst and thus changing τ. The choice of the inert solid in a fluid bed is critical because it needs to have a density and particle size distribution similar to those of the catalyst to prevent segregation in the bed. 2.4. Stripping Time. After the oil injection period, the deactivated catalyst was stripped with nitrogen to remove all volatile products. We measured the effect of stripping time at various catalyst-to-oil ratios by varying the stripping time from 300 to 800 s. In all cases there was no change in coke on catalyst beyond 600 s. Therefore, in this work we have used a stripping time of 700 s for all our runs. 2.5. Fixed-Bed Experiments. We carried out these experiments in a standard fixed-bed Pyrex reactor, 1.5 cm i.d., using the constant ts protocol with the catalyst dilution approach. We used 8.4 g of total solids, a 2.4 g/min oil injection rate, and 30 s ts. We used an oil delivery system, commonly known as a deadman, that has the ability to heat the feed close to the reaction temperature and thus minimize the initial catalyst temperature drop to less than ca. 40 K at the start of the experiment. The liquid and gas handling and analysis techniques were similar to those of our fluidbed reactor unit. We obtained coke on catalyst at the end of a run via a LECO unit.
3. Results and Discussion 3.1. Varying Time-on-Stream Protocol. Figure 1 shows changes in conversion as the catalyst-to-oil ratio is varied with corresponding changes in the catalyst time on stream. Because τ is constant, i.e., the catalyst weight and oil injection rate are constant, an increase in the catalyst-to-oil ratio implies a decrease in the amount of oil fed to the reactor and hence a decrease in
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Figure 1. Conversion as a function of the catalyst-to-oil ratio for the varying ts protocol.
Figure 3. Analysis of the individual aliquots under the variable ts protocol shown in Figure 1.
weighted difference in conversions at 7 and 8 catalystto-oil ratios as follows:
X)
Figure 2. Coke deposited on the catalyst under the varying ts protocol. Table 1. Coke Selectivity of Individual Aliquots under the Variable Time-on-Stream Protocol catalyst-to-oil ratio ts oil weight, g conversion, wt % coke yield, wt % ts oil weight, g conversion, wt % coke yield, wt % coke selectivity, wt %
Cumulative Data 8 7 6 56 64 75 1.12 1.28 1.5 76.1 75.7 74.1 4.80 4.54 4.01
5 90 1.8 72.2 3.58
4 112 2.25 69.9 3.13
3 150 3.0 65.3 2.53
Individual Aliquots 56 8 11 1.12 0.16 0.22 76.1 72.2 64.9 4.80 2.72 0.91 6.30 3.77 1.40
15 0.3 62.7 1.44 2.30
22 0.45 60.4 1.29 2.13
38 0.75 51.8 0.78 1.50
ts. Hence, for a catalyst-to-oil ratio of 8, 1.12 g of oil is injected over 56 s, whereas for a catalyst-to-oil ratio of 3, 3 g of oil is injected over 150 s. We show the actual weight of oil delivered for each experiment in Table 1. The conversion is initially weakly dependent on the catalyst-to-oil ratio; decreasing it from 8 to 5 decreases the conversion by only 4 wt %. There is a larger decrease in the conversion at very low catalyst-to-oil ratios that correspond to very long ts. Figure 2 indicates that, although the largest amount of coke is formed in the first 56 s (0.6 wt %), additional coke continues to be formed and deposited on the catalyst with increasing ts. Since τ for all experiments in Figure 1 is the same at 450 s, it is mathematically possible to determine conversions of individual aliquots or segments of oil fed. At a catalyst-to-oil ratio of 8, we add 1.12 g of oil in 56 s and obtain 76.1% conversion, whereas at a catalystto-oil ratio of 7, we add 1.28 g of oil in 64 s, obtaining 75.6% conversion. The experiment with a catalyst-tooil ratio of 7 (1.28 g of oil) may be visualized as an experiment with a catalyst-to-oil ratio of 8 (1.12 g of oil) followed by the addition of 0.16 g of oil. We calculate the conversion of this additional 0.16 g of oil from the
(1.28 × 75.6) - (1.12 × 76.1) 0.16
where X is the percent conversion of 0.16 g of oil. This value is 72.2%; therefore, although the time-averaged conversion of the catalyst-to-oil ratio experiment of 7 is 75.7, the last 0.16 g of oil added has a lower conversion. Thus, the weighted difference between any two consecutive catalyst-to-oil ratio experiments is the conversion of the additional weight of oil added to the lesser of the two ratios. We have calculated the conversion of each of the individual aliquots of oil from the data shown in Figure 1. Figure 3 shows the conversion of these individual time-averaged aliquots. Because the original data were obtained for ordinal catalyst-to-oil ratios, the size of individual aliquots varies quite significantly from 56 to 8 s. At a catalyst-to-oil ratio of 3, the time-averaged conversion was 65.3% but the conversion of the oil fed in the last 38 s had a conversion of only 52%. The decline in conversion from 76 to 52 wt % in Figure 3 is not surprising because during that time the coke deposited on the catalyst increased from 0.6 to 0.85 wt %. Because all conversions are time-weighted averages, the largest aliquot will dominate, which, in our case, is at the catalyst-to-oil ratio of 8 with ts of 56 s. Since all runs are carried out at the same τ, why does conversion vary at all? Cracking of hydrocarbons occurs predominantly on acid sites, with coke being one of the products. Coke is known to poison acid sites, making them unavailable for further participation in cracking reactions.9,10 Therefore, cracking activity and conversion are expected to decline with increasing ts because of increased coke deposition. This is indeed observed at longer ts. The first 56 s aliquot had a conversion of 76 wt % with 0.6 wt % of coke on the catalyst. As the coke on the catalyst increased to 0.85 wt %, the conversion of the last 38 s aliquot declined to 52%. Therefore, in the variable ts, constant τ protocol, conversion changes only because of catalyst deactivation. As with conversion, it is mathematically possible to determine product yields of each individual aliquot. Since coke plays such a prominent role in deactivation, we show, in Table 1, conversion and coke yields both for the experimental catalyst-to-oil ratio set and for the calculated individual aliquots. Analogous to conversion, coke yields of the individual aliquots declined more sharply than those in the experimental data. Whereas decline in conversion from 76.1 to 65.3 wt % decreased
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time-averaged coke yield from 4.8 to 2.5 wt %, the coke yield of the individual aliquots decreased from 4.8 to 0.78 wt %. Coke is expected to decrease with decreasing conversion. The question is whether the decline in coke yield is larger than that predicted by the decrease in conversion. However, to understand this point, rather than using the coke yield, i.e., the percentage of total feed converted to coke, we use coke selectivity, i.e., the percentage of total products converted to coke, which we show for the individual aliquots in Table 1. Note the sharp decline in coke selectivity between the first aliquot and the second aliquot, where depositing 0.6 wt % coke on the catalyst significantly lowers coke selectivity; this is followed by a progressive but gradual decline in coke selectivities for the subsequent aliquots. Such a change in coke selectivity caused by predepositing coke during catalytic cracking has been reported earlier.11 Overall, coke selectivity declined by a factor of 4 from the first 56 s oil aliquot to the last 38 s oil aliquot. The fact that selectivities and yields are dependent on ts is critical in catalyst evaluation. Most catalyst testing is carried out in the range of 50-80 wt % conversion, and the results regressed to obtain yields either at a constant conversion or at a constant coke yield. Under the variable ts protocol, because the deactivation of the catalyst changes, which, in turn, changes the product selectivity, regression of the results to constant conversion or coke is not meaningful. Furthermore, even if such a regression is carried out when comparing different catalysts, in a varying ts protocol, catalysts of differing activities will achieve the requisite constant conversion at different ts. Because the catalyst becomes less coke selective at longer ts, this introduces an artifact where one obtains selectivity differences at constant conversion for different catalysts solely due to differences in activity; this makes comparative evaluations via this protocol ineffective. 3.2. Constant Time-on-Stream Protocol. 3.2.1. Effects of Backmixing. Figure 4a shows activity versus space time plots at a ts of 30 s for fluidand fixed-bed reactors. Since gas oil cracks by an apparent second-order process, results from a fixed bed give a linear plot that passes through the origin using the activity function in the absence of expansion, x/(1 - x), where x is the fractional conversion.12 Using this function with results from a fluid bed gives a plot that is linear at low τ but curves at higher values. This curvature indicates that the accepted second-order activity function for plug-flow reactors is not correct for the fluid bed. The corresponding activity function for a backmixed reactor12 in the absence of expansion is x/(1 - x)2 and has been suggested previously for fluidized beds in catalytic cracking.13 As shown in Figure 4b, when we plot our fluid-bed data with this activity function, we obtain a linear response going through the origin with excellent goodness of fit, indicating that this is the appropriate activity function for a small fluidbed reactor and also implying substantial backmixing in the catalyst bed. Furthermore, at constant τ, conversion in a fluid-bed reactor is lower or equal to that in a fixed-bed reactor, depending on the extent of conversion, as it should be for a non-zero order reaction. We note that the activity functions for both fixed and fluid beds do not include an expansion factor. Gas oil cracking is an apparent rather than a true second order reaction. Individual hydrocarbon molecules react by a
Figure 4. (a) Activity vs space time relationship for fixed and fluid beds at 30 s ts. (b) Backmixed activity function vs space time for the fluid-bed reactor operated at 30 s ts.
first order process. The different molecules in gas oil have vastly different reactivities. The net effect of such differing reactivities, as shown by Weekman,14 is for the apparent overall reaction to be second order. It is possible that the pseudo second order nature of the overall reaction plays a role in making the expansion factor appear close or equal to zero. We also note that, because, within experimental error, the activity-τ plots for both types of reactors go through the origin, the inert microspheres used for blending do not contribute to catalytic activity. The slope of the activity-τ plot is the product of the global rate constant of cracking, k, and the inlet gas oil concentration, CA0. The measured rate constant is a time-averaged number because it is derived from time-averaged data. Table 2 compares yields obtained from fluid- and fixed-bed reactors at 70% conversion at the same ts of 30 s. A higher τ is required in a fluid-bed reactor to achieve the same conversion. If we use the coke formed as a gauge for catalyst deactivation, about the same number of sites are poisoned in both cases. The combination of backmixing and the lower partial pressure of hydrocarbons in the fluid bed requires higher τ to achieve the same conversion as that in a fixed-bed reactor. Reaction pathways for gas oil cracking are made up of series and parallel reactions. For a series reaction, backmixing maximizes the yield of final products, whereas plug-flow dynamics maximizes yields of intermediate products. Thus, yields from the two reactors may allow us to speculate on the overall network. Dry gas, coke, and LCO yields from the two reactors are very similar. Dry gas is mostly made as part of the initiation step, either via a noncarbocation route15 or via proton attack16 and via thermal cracking. It seems that this chemistry is not affected by bed dynamics. Ho17 and John and Wojciechowski18 have noted the complex nature of coke formation via primary and secondary pathways. Recently, den Hollander et al.19
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Table 2. Comparison of Yields at 70% Conversion between the Two Reactors at 30 s Time on Stream yield (wt %)
hydrogen methane ethane ethylene propane propylene n-butane i-butane n-butenes isobutylene LPG olefins LPG paraffins LPG olefins, % dry gas LPG gasoline coke total coke, mg LCO HCO
fluid-bed reactor (τ ) 100 s)
fixed-bed reactor (τ ) 70 s)
0.08 0.75 0.53 0.81 0.92 5.16 0.79 3.72 3.98 1.52 10.7 5.4 66 2.2 16.1 45.8 5.9 83 18.5 11.5
0.16 0.94 0.79 0.82 1.16 4.39 0.79 3.64 2.90 1.02 8.3 5.6 60 2.7 13.9 47.3 6.1 85 18.4 11.7
showed in a laboratory riser unit that coke was formed only via a primary reaction of the gas oil and not from secondary reactions. In our case, coke yields, at the same conversion, are very similar in fixed- and fluid-bed reactors, indicating that backmixing does not influence the coke yield. This suggests that coke formation is predominantly primary, or we speculate that the kinetics of coke formation are zero order in hydrocarbon concentration. The main difference in the products from the two reactors is in the gasoline-LPG split and the high olefinic nature of the LPG from the fluid-bed reactor. The latter gives 1.5-2% absolute lower gasoline and higher LPG yields than the fixed-bed reactor. The increase in LPG is mainly due to the increase in olefins, whereas the paraffins are essentially the same. These observations are as expected from the sequential nature of the reactions and the effect of backmixing on such reactions. In an ideal continuous stirred tank reactor, the concentration of reactants and products in the reactor are the same as that at the exit. In our fluidbed reactor, even though it does not operate in a truly continuous fashion, the olefin concentration in the reactor will be higher at any point in time than at an equivalent stage in a plug-flow fixed-bed reactor. Hence, β scission of a larger gasoline range olefin would give two smaller olefins, therefore increasing LPG olefins and decreasing the gasoline yield. The LCO yields between the two reactors are the same. Since our feed has very low levels of material in the LCO fraction, our measured LCO product is that converted from the gas oil. However, as defined, conversion does not take LCO formation into account. Because the space time we use is different, to obtain equivalent conversion of gas oil in the two reactors, we cannot comment on whether the sequential nature of the reactions affects LCO formation in the backmixed reactor. 3.2.2. Effect of Different Times on Stream. We carried out a series of runs at 15, 30, 60, and 90 s ts, keeping all other experimental parameters constant. Since the oil injection rate is constant, the total amount of oil injected to the reactor increases monotonically with 0.7 g of oil being injected in 15 s ts and 4.2 g of oil
Figure 5. Backmixed activity function vs space time in a fluidbed reactor for 15, 30, 60, and 90 s ts (rate constant kCAO of cracking in parentheses).
Figure 6. Coke yield vs backmixed activity function in a fluidbed reactor at 15, 30, 60, and 90 s ts.
in 90 s ts. The backmixed activity-τ plots in Figure 5 show that the slope of the plot decreases with increasing ts. The slope is equal to the product of k with CA0, and because CA0 is the same for all runs, the decrease in the slope of the lines is due to a decrease in the global rate constant. Though the time-averaged rate constant of cracking declines by ca. 3.5 times as ts is increased from 15 to 90 s, the instantaneous rate constant of cracking declines much more rapidly. Since the primary cause for deactivation with ts is the deposition of coke, we examine coke yields at different ts’s (Figure 6). Typically, in fixed-bed reactors, the coke yield-activity plot intersects the coke ordinate at either the CCR value of the gas oil feed or some fraction thereof.20 Here, even though the CCR of our feed is very low, we observed a significantly higher positive intercept for all four sets, with a much larger coke yield at 15 s ts compared to the other three ts’s. At 70% conversion, backmixed activity ) 7.78, the 15 s ts set shows a coke yield of 8.3%, whereas with the 30 s ts set, we obtain a 30% lower coke yield of 5.8 wt %, with a further increase in ts having only a small impact on the coke yield. The very high coke yield we report at very short ts, where the least amount of oil is delivered, is because, as shown previously, the largest amount of coke forms early in the oil injection cycle. Since coke yield is defined on feed basis and since the amount of oil fed to the reactor varies for different times on stream, we look at the total absolute coke deposited on the catalyst as a function of τ (Figure 7). We have also included results from our fixed-bed reactor operated at a ts of 30 s. Even though coke yield decreases with increasing ts, the absolute total coke deposited on the catalyst, as expected, increases with ts. Coke deposited on the catalyst may be divided into two regions: one is the apparent intercept, and the second is the region with increasing space times. Since the τ-activity plot is
Ind. Eng. Chem. Res., Vol. 42, No. 3, 2003 431 Table 4. Yields at 70% Conversion in a Fluid-Bed Unit at Different Times on Stream yield (wt %)
Figure 7. Total coke deposited on the catalyst in a fluid-bed reactor at 15, 30, 60, and 90 s ts and in a fixed-bed reactor at 30 s ts. Table 3. Grams of Coke per Gram of Catalyst at 70% Conversion for Different Times on Stream t s, s
τ for 70% conversion, s
coke/catalyst, mg/g
total coke, mg
kCAO
15 30 60 90
70 100 180 270
17.7 17.6 17.6 18.0
58 82 148 227
0.108 0.074 0.045 0.030
linear and passes through the origin, coke deposited beyond the intercept is directly proportional to activity. This is reasonable because an increase in activity should linearly increase the coke deposited on the catalyst. We observe this in the fixed-bed unit; the difference in the fluid bed is the apparent intercept. The intercept at τ equal to zero should only reflect the 3.1 mg of CCR deposited on the catalyst in 30 s ts. The significantly larger intercept is due to the unique nature of gas oil contacting in a fluidized bed. The initial amount of oil contacts the entire bed of clean catalyst, which has a propensity to make predominantly coke. The large intercept is due to the rapid initial deposition of coke on the entire catalyst. In a fixed-bed reactor, where the oil-catalyst contacting is sequential, only the initial amount of oil comes into contact with clean catalyst on top of the fixed bed. This would explain the different coke-activity behavior in fixed- and fluid-bed reactors. At 70% conversion, the total coke deposited increases from 58 to 227 mg as ts increases from 15 to 90 s. The 3.9 times increase in the total amount of coke formed at the longer ts corresponds exactly with the 3.9 times higher τ required to achieve the same conversion and with the decline in the overall rate constant of cracking by a factor of 3.6 (Figure 5 and Table 3). Therefore, even though the milligrams of coke deposited per gram of active catalyst is constant at ca. 18 (Table 3), this cannot be used as the sole measure of deactivation because more sites are required at longer ts to achieve the same conversion. The importance of ts in catalyst deactivation has been well established.21-23 Since the rate of coke deposition is higher at the start of our experimental run, the critical effect of ts, in our case, may be due to (i) coke forming intermediates that are measured in the experiment as coke but transform over time to coke that deactivates the necessary catalytic sites or (ii) deactivation that depends on the distribution of acid sites, with the sites contributing to higher turnover being deactivated first. Both of these models would result in a larger number of sites being needed at longer ts to obtain the required conversion.
hydrogen methane ethane ethylene propane propylene n-butane i-butane n-butenes isobutylene LPG olefins LPG paraffins LPG olefins, % total dry gas LPG gasoline coke LCO HCO LCO/HCO
ts ) 15 s τ ) 70 s
ts ) 30 s τ ) 100 s
ts ) 60 s τ ) 180 s
ts ) 90 s τ ) 270 s
0.09 0.81 0.50 0.90 0.85 5.39 0.75 3.66 4.07 1.65 11.1 5.3 68 2.3 16.4 43.1 8.26 17.5 12.5 1.41
0.08 0.75 0.53 0.81 0.92 5.16 0.79 3.72 3.98 1.52 10.7 5.4 66 2.1 16.1 45.8 5.88 18.5 11.5 1.60
0.08 0.68 0.55 0.70 1.02 5.03 0.93 4.39 3.99 1.36 10.4 6.3 62 2.0 16.7 45.9 5.30 18.8 11.2 1.69
0.08 0.75 0.62 0.68 1.27 4.48 1.01 4.42 3.32 1.08 8.9 6.7 57 2.1 15.6 46.9 5.41 19.4 10.7 1.82
As ts increases by a factor of 6, i.e., from 15 to 90 s, the absolute coke deposited increases by a factor of 3.9. Hence, the coke yield defined on a feed basis decreases with increasing ts (Figure 6). This impacts all other yields since all product yields are defined on a feed basis. Table 4 shows product yields at 70 wt % conversion for the four different ts sets. An increase in ts from 15 to 90 s results in a 2.85 wt% lower coke yield, a 3.8 wt% higher gasoline yield, and a 0.8 wt% lower LPG yield. At longer ts the gasoline yield increase is not only due to a decrease in the coke yield but also due to the decrease in LPG. We propose that at long ts’s the deactivation of sites, as noted above, mitigates the cracking of gasoline range hydrocarbons to LPG. Furthermore, the LPG is less olefinic with increasing ts. This may be due to the factor of 4 higher τ needed at ts ) 90 s to achieve the same 70% conversion for the run at ts ) 15 s. The total amount of coke is also much higher at ts ) 90 s. We cannot rule out the possibility that this coke plays a role in the hydrogenation of olefins.24 The increase in LCO as ts increases is due to higher space times needed to obtain an equivalent conversion of gas oil. 3.3. Comparison of the Results from Fluid- and Fixed-Bed Reactors. Earlier we showed that at a constant ts of 30 s the fixed- and fluid-bed reactors displayed the expected characteristics of a plug-flow and backmixed reactor with respect to both conversion and selectivities (Table 2). However, at longer ts of 90 s, the gasoline and LPG yields from the fluid-bed reactor approach those obtained in a fixed-bed reactor at 30 s ts (Table 5). Even more striking is the similarity in the LPG olefin content. This similarity is fortuitous and comes about for different reasons. In fixed-bed reactors, which by definition maximize yields of intermediate products compared to backmixed systems, less gasoline range olefins are cracked to LPG range olefins, thus giving a lower olefin content in the LPG. At the same ts backmixing results in lower gasoline yield and higher olefinic LPG. At longer ts the LPG olefin content decreases because of increased catalyst deactivation due to the higher amount of coke made (Table 3). Deactivation results in the poisoning
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Table 5. Comparison of Product Yields (wt%) at 70% Conversion from a Fluid-Bed Reactor at 90 s Time on Stream with Those from a Fixed-Bed Reactor at 30 s Time on Stream
H2 methane ethane ethylene propane propylene n-butane i-butane n-butenes isobutylene LPG olefins LPG paraffins LPG olefins, % total dry gas LPG gasoline coke LCO HCO LCO/HCO
fluid-bed reactor at 30 s ts τ ) 100 s
fixed-bed reactor at 30 s ts τ ) 70 s
fluid-bed reactor at 90 s ts τ ) 270 s
0.08 0.75 0.53 0.81 0.92 5.16 0.79 3.72 3.98 1.52 10.7 5.4 66 2.2 16.1 45.8 5.9 18.5 11.5 1.6
0.16 0.94 0.79 0.82 1.16 4.39 0.79 3.64 2.90 1.02 8.3 5.6 60 2.7 13.9 47.3 6.1 18.4 11.7 1.6
0.08 0.75 0.62 0.68 1.27 4.48 1.01 4.42 3.32 1.08 8.9 6.7 57 2.1 15.6 46.9 5.4 19.4 10.7 1.8
of acid sites that crack gasoline range olefins to LPG range olefins. Also, as noted earlier, the higher amount of coke present may play a role in olefin hydrogenation. These combined effects in the fluid bed, which are the result of increased deactivation, make the product distribution from a fluid-bed reactor appear similar to that from a fixed-bed reactor. The higher LCO yield in the fluid-bed reactor is due to the higher 270 s τ required to achieve 70% conversion at 90 s ts. Because as noted earlier LCO formation is not included in our definition of conversion, the pathway heavy cycle oil (HCO) to LCO may be considered as a separate reaction, with the formation of LCO for this reaction increasing, as expected, with increasing τ. 4. Concluding Remarks As noted by Wallenstien et al.8 and as shown by us, conversion in fluid-bed reactors operated under the variable ts protocol is weakly dependent on the catalystto-oil ratio. The variable ts protocol operates at constant τ. Any change in conversion is solely due to deactivation. If the choice of catalyst and/or gas oil feed is such that deactivation is minimized, there will be very little variation in conversion as catalyst-to-oil ratios are changed. In the extreme case of severely hydrotreated feed that has a much lower propensity for coke formation and deactivation of the catalyst, we should expect no change in conversion. Testing while varying ts is not unique to fluid beds. Many research groups have historically practiced such an approach with fixed-bed reactors.2 And, it has been previously noted4 that such an approach “ensures that selectivities are determined at different feed reaction times if catalysts of different activities are investigated”, thus giving artificial differences in product yields. Our work has identified the artifact that causes this problem. We have shown that, at very short ts, coke yields are very high. Once a certain amount of coke is deposited on the catalyst, selectivity toward additional coke formation decreases substantially. This directly impacts other product yields because yields for cracking
catalysts are compared at constant conversion. An objective of such comparisons is often to develop or choose a catalyst that has a lower coke yield and higher gasoline and/or LPG yield. A higher activity catalyst will attain the desired conversion at longer ts. Since the additional oil being fed to the reactor is preferentially converted to products other than coke, the net effect is for the higher activity catalyst to show a lower coke yield. Therefore, it is essential in order to obtain correct evaluations that one varies conversion not by varying ts but by varying space time at constant ts. For catalytic cracking, an appropriate way to change space time is by changing catalyst mass. Changing catalyst mass has an undesirable side effect of changing the hydrodynamics of the bed. Although important for both types of reactors, hydrodynamics is expected to play a greater role in fluid beds. The proper way to circumvent this problem is by using the catalyst dilution approach as described in this work and by McElhiney.7 By using the constant ts protocol, we have shown that small fluid beds are significantly backmixed. This is obvious from our comparison with a fixed-bed reactor with experiments carried out at the same ts. However, by increasing ts, we show that yields from a backmixed fluid-bed reactor can be made to approach those from a plug-flow fixed-bed reactor. The results from the fluid bed are in large part due to the increased deactivation and the necessity therefore of using higher τ to get the required conversion. Although the similarities were striking for our combination of catalyst and gas oil (Table 5), we believe the general trend will hold true. In FCC catalyst evaluation literature, conversion variation in a laboratory reactor is frequently related to changes in the “catalyst-to-oil ratio”. Although it is the correct parameter to describe the operation of a commercial riser or a circulating pilot-plant riser reactor, it has no significance for describing a laboratory fixed-bed or fluid-bed reactor. The operation of laboratory reactors and the conversion obtained are uniquely defined by space time at a certain time on stream. Acknowledgment The authors gratefully acknowledge Harshad Patel for his assistance in the laboratory and Engelhard Corp. for permission to publish the work. Literature Cited (1) Krishna, A. S.; Arndt, J. H.; Kuehler, C. W.; Kramer, D. C. Refiner details ‘best practices’ approach to catalyst selection. Oil Gas J. 1996, 14 (Oct), 44. (2) Moorehead, E. L.; McLean, J. B.; Cronkright, W. A. Microactivity evaluation of FCC catalysts in the laboratory: principles, approaches and applications. Stud. Surf. Sci. Catal. 1993, 76, 223. (3) Wallenstien, D.; Alkemade, U. Modeling of selectivity data obtained from microactivity testing of FCC catalysts. Appl. Catal. 1996, 137, 37. (4) Wallenstien, D.; Harding, R. H.; Witzler, J.; Zhao, X. Rational assessment of FCC catalyst performance by utilization of micro-activity testing. Appl. Catal. 1998, 167, 141. (5) Kayser, J. C. (Kayser Technologies, Inc.). Versatile fluidized bed reactor. U.S. Patent 6,069,012, 2000. (6) Stockwell, D. M.; Wieland, W. S.; Himpsl, F. L. Catalyst evaluation using fixed fluidized bed reactors: protocol is critical to correct performance ranking. Int. Conf. Ref. Proc. AIChE 1998, 237. (7) McElhiney, G. FCC catalyst selectivity determined from microactivity tests. Oil Gas J. 1988, 8 (Feb), 35.
Ind. Eng. Chem. Res., Vol. 42, No. 3, 2003 433 (8) Wallenstien, D.; Haas, A.; Harding, R. H. Latest developments in microactivity testing: influence of operational parameters on the performance of FCC catalysts. Appl. Catal. 2000, 203, 23. (9) Derouane, E. G. Factors affecting the deactivation of zeolites by cracking. Stud. Surf. Sci. Catal. 1985, 20, 221. (10) Hopkins, P. D.; Miller, J. T.; Meyers, B. L.; Ray, G. J.; Roginski, R. T.; Kuehne, M. A.; Kung, H. H. Acidity and cracking activity changes during coke deactivation of ultrastable Y zeolite. Appl. Catal. 1996, 136, 29. (11) Cody, I. A.; Stuntz, G. F.; McKnight, W. G. (Exxon Research & Engineering Co.). Hydrocarbon catalytic cracking utilizing a precoked catalyst. U.S. Patent 5,158,670, 1992. (12) Levenspiel, O. Chemical Reaction Engineering; Wiley: New York, 1964; p 116. (13) Blanding, F. H. Reaction rates in catalytic cracking of petroleum. Ind. Eng. Chem. 1953, 1186. (14) Weekman, V. W. A model of catalytic cracking conversion in fixed, moving and fluid-bed reactors. Ind. Eng. Chem. 1968, 7, 90. (15) McVicker, G. B.; Kramer, G. M.; Ziemiak, J. J. Conversion of Isobutane over Solid AcidssA Sensitive Mechanistic Probe Reaction. J. Catal. 1983, 83, 286. (16) Haag, W. O.; Dessau, R. M. Duality of the Mechanism for Acid-Catalyzed Paraffin Cracking. Proceedings of the 8th International Congress on Catalysis, Berlin, 1984. (17) Ho, T. C. Study of Coke formation in Resid Catalytic Cracking. Ind. Eng. Chem. Res. 1992, 31, 2281.
(18) John, T. M.; Wojciechowski, B. W. On identifying the primary and secondary products of the catalytic cracking of neutral distillates. J. Catal. 1975, 37, 240. (19) den Hollander, M. A.; Makkee, M.; Moulijn, J. A. Coke formation in fluid catalytic cracking studied with the microriser. Catal. Today 1998, 46, 27. (20) Mott, R. W. New concept measures catalyst performance. Oil Gas J. 1987, 26 (Jan), 73. (21) Pachovsky, R. A.; Best, D. A.; Wojciechowski, B. W. Applications of the time-on-stream theory. Ind. Eng. Chem. Process Des. Dev. 1973, 12, 254. (22) Grzesik, M.; Skrzypek, J.; Wojciechowski, B. W. The catalyst decay behavior in fluiduized-bed reactors using time-onstream theory. Chem. Eng. Sci. 1990, 45, 267. (23) Rice, N. M.; Wojciechowski, B. W. Catalyst decay in the presence of chain processes in catalytic cracking. Can. J. Chem. Eng. 1991, 69, 1100. (24) Wojciechowski, B. W.; Corma, A. Catalytic Cracking: Catalysts, Chemistry and Kinetics; Marcel Dekker: New York, 1986; p 200.
Received for review August 13, 2002 Revised manuscript received November 4, 2002 Accepted November 4, 2002 IE0206251
