Latest developments in microactivity testing: influence of operational parameters on the performance of FCC catalysts

Applied Catalysis A: General 203 (2000) 23–36

Latest developments in microactivity testing: influence of operational parameters on the performance of FCC catalysts D. Wallenstein a,∗ , A. Haas a , R.H. Harding b a

b

Grace GmbH, Postfach 1445, 67545 Worms, Germany W.R. Grace and Co.-Conn. 7500 Grace Drive, Columbia, MD 21044, USA

Received 2 August 1999; received in revised form 25 January 2000; accepted 25 January 2000

Abstract An accurate assessment of catalyst performance is the most important goal in the testing of fluid catalytic cracking (FCC) catalysts since the catalyst is a key contributor to the overall unit performance. For this purpose small scale tests in the laboratory have been the workhorse in the industry because they are less expensive and time-consuming to operate than circulating riser pilot units. Two reactor types are common in microactivity testing, simple fixed bed (plug-flow) reactors and fixed fluidised bed reactors. Pitfalls have been identified for both experimental modes and this paper discusses the strengths and weaknesses of these two techniques. This work demonstrates that catalyst testing in small-scale fixed fluidised bed reactors can result in erroneous catalyst ranking, while the use of improved fixed bed reactors has a better agreement with riser pilot units. Furthermore, reactors of either type that employ in-situ regeneration result in unrealistically high coke and hydrogen yields due to the oxidation of contaminant metals, and therefore advanced deactivation procedures for metallated FCC catalysts cannot be utilised in small-scale tests with this technique. © 2000 Elsevier Science B.V. All rights reserved. Keywords: Fluid catalytic cracking; Microactivity tests; Fixed bed reactors; Fixed fluidised bed reactors

1. Introduction Fluid catalytic cracking is a pivotal process in the modern refinery used to upgrade vacuum gas-oils and resids from atmospheric and vacuum distillation [1]. Small changes in catalyst performance have a substantial impact on FCC unit profitability and therefore a realistic catalyst evaluation in the laboratory is of utmost importance. It is widely accepted that circulating riser pilot units provide the best small-scale simulation of commercial operations, however, these units require both high investment and operating costs. Thus, microactivity tests are still the most commonly ∗ Corresponding author. Tel.: +49-6241-4033-12; fax: +49-6241-4039-0368.

used methods to characterise the performance of FCC catalysts. The well-known microactivity test (MAT) ASTM-D 3907-92 has been described as a tool for selectivity testing [2,3], and has been used widely in the refining industry. However, numerous pitfalls have been identified for this approach, which often result in an erroneous catalyst evaluation [4–11]. Recent efforts have focused on improving the results from microactivity testing. Researchers have used two basic strategies to improve the test. The first approach has been to modify and improve the fixed-bed ASTM method with regard to catalyst bed geometry, contact-time, feed pre-heat, reaction temperature and catalyst-to-oil adjustment. The second approach, which is a more substantial change in methodology, has been to convert the MAT reactor

0926-860X/00/$ – see front matter © 2000 Elsevier Science B.V. All rights reserved. PII: S 0 9 2 6 - 8 6 0 X ( 0 0 ) 0 0 4 6 7 - 1

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D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

from a fixed bed to a fixed fluidised bed (FFB) reactor design. Historically, the use of fluidised beds for microactivity testing increased in order to provide a better simulation of FCC units with fluidised catalyst beds. However, with the advent of commercial riser reactors, the use of the FFB technique in laboratory catalyst testing decreased due to (i) backmixing in these reactors, (ii) very long catalyst-time-on-stream necessary in this reactor-type to achieve a reasonable fluidisation, and (iii) better simulation of riser conditions by plug-flow reactors [4,5]. Many laboratories have switched back to the traditional fixed bed for catalyst screening and to riser pilot units for a final catalyst evaluation. However, in spite of the challenges described above, the FFB approach has recently experienced a renaissance, mainly due to ease of automation [12–14]. Another interesting development in the ongoing improvement of microactivity testing has been the use of in-situ regeneration of the coked catalyst following the cracking experiments. This technique, which can be used in either fixed bed or fluidised bed testing, is primarily introduced in order to improve testing efficiency and ease of automation. This paper focuses on assessing the strengths and weaknesses of these attempts to improve the standard microactivity test. Experiments were performed in both fixed and fluidised bed equipment to contrast the results from the two approaches. The data shows that an improved fixed bed has fewer limitations than the use of a fluidised bed for accurate assessment of catalyst performance.

2. Experimental The methods employed are thoroughly described in [9,10,15–18]. Only the relevant details for this work are given in this section. 2.1. Fixed bed testing The fixed bed results were obtained with the short contact-time (SCT) microactivity test described in the [9,10] references. Briefly, the SCT-MAT is a fixed bed with an annular bed design. The thickness of the catalyst layer between reactor wall and reactor core was 2.5 mm. Catalyst-to-oil ratio was varied by changing

the mass of catalyst while the total amount of feed and time-on-stream were kept constant at 1.5 g and 12 s, respectively. The volume of the catalyst bed was 10 ml at each catalyst-to-oil ratio and this was achieved by diluting the catalyst with glass beads of 0.2–0.3 mm diameter. For each cracking experiment a new portion of catalyst and glass beads were used. Before the cracking experiment was started the reactor was equilibrated for 45 min in flowing nitrogen at the reaction temperature of 833 K. The feed was pre-heated prior to injection into the catalyst bed to about 630 K. After passing through the catalyst bed the liquid products were collected in a receiver. The gaseous products passed through the liquid products and were collected in a vessel over water. The catalyst bed and liquid products were then stripped with a stream of 30 cc/min nitrogen to a total gas volume (crack gas+nitrogen) of 600 ml which was determined by the amount of displaced water. Coke on catalyst was determined by a carbon analyser type CS-344 [15]. 2.2. Fixed fluidised bed testing The fluidised bed results were obtained with an ACE® unit from Xytel and Kayser Technologies. The ACE unit is an automated fluidised bed test. The reactor is fluidised with a stream of 100 cc/min nitrogen. In these studies, 9 g of catalyst were fluidised and stabilised at the reaction temperature of 794 K. Feed was flowed over the catalyst at a rate of 1.2 g/min. Time-on-stream was varied from 50–150 s to change the catalyst-to-oil ratio. After reaction, the catalyst was stripped by nitrogen for an additional 500 s to remove adsorbed reaction products. Coke on catalyst was determined by in-situ regeneration with fluidised air by heating to 966 K. 2.3. Analysis of products, data evaluation The liquid and gaseous products obtained by the SCT-MAT and the ACE unit were analysed by the gaschromatographic techniques referenced in [15]. The yields were calculated as weight percent of the feed. The fractions of gasoline, light cycle-oil and heavy cycle-oil are determined by the cut points at 489 and 611 K. Conversion is defined as 100 (wt.%)−light cycle-oil (wt.%)−heavy cycle-oil (wt.%).

D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

Curve fitting to conversion and yield data and interpolation of yields at constant conversion were performed with the method described in [16]. 2.4. Catalyst pre-treatments Catalysts were deactivated using two methods. Catalysts which were tested in the metals-free condition were pre-treated with 100% steam in a fluidised bed at atmospheric pressure at 1089 K for 5 h [15]. Deactivation of FCC catalysts, in the presence of vanadium and nickel was performed by metallation of the catalysts with the incipient wetness method followed by cyclic propylene steaming described in [17]. Briefly, the impregnation was carried out using a solution of vanadyl naphthenate and nickel naphthenate in toluene. The volume of the impregnation solution was ca. 50% higher than the pore volume of the catalyst. The excess toluene was evaporated under vacuum at 403 K. The impregnated catalysts were then heated in air at 927 K for 3 h to decompose the naphthenates and also to burn off any coke. Thereafter the catalysts were exposed to 30 cycles of oxidising and reducing gases at 1061 K. One redox cycle consists of: 50% nitrogen + 50% steam, 50%, 5% propylene in nitrogen + 50% steam, 50% nitrogen + 50% steam, 50%, 4000 ppm SO2 in air + 50% steam, whereby the length of each step is 10 min and the 30th cycle ends in reducing atmosphere, i.e. 10 min with 50%, 5% propylene in nitrogen +50% steam. Thereafter the catalysts were cooled down to room temperature under nitrogen flow.

25

• the large gas-volume formed during the cracking experiment entails high gas velocities in the reactor and both can generate channels in the catalyst bed, • the endothermic cracking reactions and the heat consumed by the evaporation of the feed results in large temperature drops dependant on the activity of the catalyst, • coke profiles. In order to overcome these limitations, an improved fixed bed design has been developed in the form of a diluted annular bed illustrated in Fig. 1 (Reactor B). In this reactor configuration the catalyst bed is arranged around a glass core and diluted with glass-beads. Thus, channelling effects are minimised by the longer catalyst bed and the dilution with a heavier, inert material. Moreover, the dilution with glass beads enhances the interparticle space in the catalyst bed and thus provides more room for gas expansion which also reduces the formation of channels. Temperature measurements in the catalyst bed were performed during the cracking experiments. The results are depicted in Fig. 2 which also contains the temperature profile in an FFB reactor. The temperature profile of the diluted annular bed displays a significantly improved isothermicity compared with the dense cylindrical bed and this attributed to the thermal mass of the glass core and the diluting agent. The FFB data are discussed at the end of this section. Following the cracking experiment several catalyst layers of the dense cylindrical and diluted annular fixed bed were analysed for coke content and the results are given in Table 1. These data and the visual appearance of the different catalyst beds after cracking experiments, revealed a virtually homogeneous distribution of coke in the diluted annular catalyst bed whereas the ASTM design exhibited a distinct coke profile. Coke can be Table 1 Coke profiles in dense and diluted fixed catalyst bedsa

3. Factors that influence catalyst performance

Catalyst bed

3.1. Catalyst bed geometry The ASTM-MAT technique was developed on the basis of a fixed bed reactor realised by a cylindrical, undiluted, dense catalyst bed (Fig. 1, Reactor A). The following drawbacks have been identified for this design:

Dense

Diluted

Coke concentration on catalyst (wt.%) Position Top Middle Bottom

0.42 0.51 0.55

0.51 0.54 0.53

a Catalyst properties: Table 4, catalyst D; feed properties: Table 2.

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D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

Fig. 1. Different reactor types for testing of FCC catalysts.

formed by several reaction pathways and the results obtained by the diluted annular bed indicate similar reaction kinetics for coke formation in all layers in this set-up. The homogeneous colour and similar coke concentrations imply that some of the feed molecules, responsible for coke formation by primary reactions, have their first catalyst contact at the top layer and others at deeper layers. Moreover, similar rates for coke formation by secondary and tertiary reactions (hydrogen transfer, dehydrogenation, polymerisation) over the whole catalyst bed can be concluded from these findings. Another method to overcome the limitations of the undiluted, dense catalyst bed is the use of a FFB (Fig. 1, Reactor C). Fig. 2 clearly reveals the FFB reactor to have the smoothest temperature drop. This, however, can be partially ascribed to the much longer oil injection time of 90 s compared with the 12 s for the fixed beds. Thus, this reactor provides somewhat better isothermicity and better feed dispersion than the diluted annular fixed bed.

3.2. Feed pre-heat In the standard ASTM MAT technique, the feed is pre-heated to the reactor temperature by heating the oil injection capillary by a temperature controlled resistance heater. However, in commercial FCC units, the feed is pre-heated to moderate temperatures, typically between 420–650 K whereas the catalyst in the transfer line between regenerator and riser has a temperature of about 920 K. Therefore, the feed is exposed to a thermal shock on first contact with the catalyst. Nevertheless, many laboratories adjust the feed temperature to the reaction temperature. In order to investigate the influence of feed pre-heat on catalyst performance the following MAT experiments were performed with the fixed bed apparatus: In the first experiment the oil was heated to the reactor temperature prior to injection into the catalyst bed, whereas in the second experiment the feed temperature was about 200 K lower, hence, relatively cold

D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

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Fig. 2. Temperature profiles in different reactor types.

liquid feed was injected onto the hot catalyst bed. Further details on the experiments are given in [9,16] and the feed properties are compiled in Table 2. Two catalysts were tested in each protocol and the findings together with the respective selectivities measured in the Devison Circulating Riser (DCR) Pilot Unit [18] are presented in Table 3. The DCR is widely accepted by the refinery industry as the most reliable simulation of commercial FCC units. Conversions and selectivities observed in the DCR and with the low feed pre-heat in the MAT showed good agreement for both, absolute yield values and catalyst ranking. Pre-heating the feed to the reaction temperature reduced the relative conversion difference between the catalysts from 6 to 3%. Moreover catalyst B produced significantly more coke and hydrogen than catalyst A whereas the data generated in the DCR and the MAT with the lower feed pre-heat demonstrated a reverse ranking for these selectivities and a better agreement for the conversion difference. Clearly, more realistic catalyst ranking is achieved without pre-heating the feed to the reactor temperature. We believe that the FFB approach could also benefit from lower feed pre-heat. However, the feed is

Table 2 Feedstock properties Vacuum gas-oil Classification API gravity at 289 K Aniline point (K) Sulphur (wt.%) Total nitrogen (wt.%) Basic nitrogen (wt.%) Conradson carbon (wt.%) Average molecular weight

Paraffinic 23.9 365 0.73 0.10 0.042 0.33 388

Caromatic (%) Cparaffinic (%) Cnaphthenic (%)

19.6 59.8 20.6

D-1160 simulated distillation (vol.%) IBP 5 20 40 60 80 95 FBP

513 K 584 K 640 K 684 K 724 K 770 K 820 K 895 K

K factor

11.81

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D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

Table 3 Influence of feed pre-heat on FCC catalyst rankinga Method Feed pre-heat

DCR Low

Microactivity test Low

High

Catalyst A (+)

B (+)

A (*)

B (*)

A (*)

B (*)

Activity at constant catalyst-to-oil ratio Conversion (%) 73

76

70

74

70

72

Yields at constant conversion of 68% (wt.% f f) Hydrogen 0.042 C1 +C2 2.9 Propylene 4.2 Propane 0.81 i-Butylene 2.5 i-Butane 1.7 C3 Olefinicityb 83.3 C4 Olefinicityb 76.4 LPG 14.0 Gasoline 49.0 Light cycle oil 21.2 Heavy cycle oil 10.8 Coke 2.2

0.034 2.5 4.2 0.75 1.7 2.3 84.7 67.3 13.5 49.7 21.4 10.6 2.2

0.13 1.8 4.4 0.66 2.7 1.9 87.2 76.1 15.5 48.3 21.9 10.1 2.3

0.11 1.4 4.2 0.60 1.9 2.4 87.5 68.6 14.5 49.8 22.0 10.0 2.1

0.14 2.1 4.4 0.84 2.5 2.0 86.2 73.7 15.2 47.7 22.0 10.0 2.8

0.21 1.8 4.1 0.74 1.9 2.1 83.0 70.4 14.1 48.5 22.3 9.7 3.4

Gasoline composition (%) n-Paraffins i-Paraffins Aromatics Naphthenes Olefins

3.4 24.7 29.6 11.9 30.7

3.9 23.1 25.6 9.3 38.1

4.0 28.0 25.9 9.9 32.1

4.3 23.9 26.5 9.4 35.9

4.3 27.1 27.0 10.2 31.5

3.1 20.5 27.4 10.5 39.7

Catalyst properties

Catalyst A

Catalyst B

Deactivation Al2 O3 (wt.%) RE2 O3 (wt.%) Zeolite surface area (m2 /g) Matrix surface area (m2 /g) Unit-cell size (Å)

Hydrothermal: 5 h, 41.5 0 180 34 24.23

1089 K, 100% steam, 1 atm, fluidised bed 43.8 2.2 72 42 24.30

a b

Feed properties: Table 2; curve fitting of activity and yield data was performed with (+) 4 observations, (*) 10 observations per catalyst. C3 and C4 composition in %.

pre-heated to the reaction temperature in the FFB results reported in this work. 3.3. Adjustment of reaction severity The latest developments in fixed bed and fixed, fluidised bed microscale testing have been motivated by the idea of an in-situ regeneration of the catalyst to enable an unattended unit operation and thus to enhance productivity and efficiency. The pitfall of this technique is the adjustment of reaction severity which cannot be performed by variation of catalyst mass

as is the case in each commercial riser. McElhiney claimed an identical catalyst performance for the three common catalyst-to-oil adjustment procedures such as variation of (i) feed rate, (ii) feed injection time and (iii) catalyst mass [2]. These experiments were performed with an ASTM-type MAT using 30 s catalyst time-on-stream. However, more recent detailed studies using the SCT-MAT described in [9,10] have revealed a significant influence of the three approaches on catalyst ranking. The negative effects of variation of feed injection time can readily be deduced from the discussion on ‘catalyst time-on-stream’ in Section 4.2.

D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

The consequences of varying the feed rate on catalyst performance is described and exemplified in [9,10] in great detail. Due to the relevance of this issue for the present paper a brief discussion is given in the following: In most catalyst evaluations, product slates are compared at constant conversion. In cases where catalysts have different activities the catalyst-to-oil ratio is different at constant conversion and the variation of reaction severity by feed rate at constant catalyst mass results in the so-called ‘severity effects’ described below. 3.3.1. Temperature effect High activity catalysts need a low catalyst-to-oil ratio to achieve a certain conversion which leads to a higher temperature drop in the catalyst bed due to both the stronger cooling effect of the larger feed mass on the catalyst bed and the conversion of the larger feed mass by endothermic cracking reactions. Consequently, the ratio of thermal reactions to catalytic cracking reactions is reduced. Furthermore, hydrogen transfer reactions which have a lower activation energy than cracking reactions and are also exothermic, are promoted at lower catalyst-to-oil ratios. 3.3.2. Concentration effect A low catalyst-to-oil ratio corresponds with a large feed mass. Consequently the concentration of the feed molecules in the reactor bed of a high activity catalyst is higher than in the reactor bed of a low activity catalyst. The higher concentration favours bimolecular reactions such as hydrogen transfer. A study was conducted investigating these effects on catalyst performance. Two catalysts which only differed in the degree of rare-earth exchange were tested in the DCR, Grace Davison’s SCT-MAT using (i) feed rate and (ii) catalyst mass for the adjustment of reaction severity and an ASTM-type MAT with feed rate variation. More experimental details are described in [9,16,18]. The characterisation data of the paraffinic vacuum gas-oil used for the experiments are given in Table 2. The samples were deactivated at 1089 K for 5 h in 100% steam prior to microactivity testing. Catalyst properties and selectivity data are summarised in Table 4 and show the following: catalyst ranking and differentiation agreed very well for data

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measured by the DCR and the SCT-MAT using catalyst mass for the variation of reaction severity. By contrast, the ASTM-type and SCT-microactivity tests with reaction severity adjusted by feed rate, gave a more pronounced differentiation for gasoline, LPG, and C4 -olefinicity compared with the other two methods. Moreover, a reversed ranking for bottoms cracking was observed. Bottoms cracking increased at higher rare-earths content which is not expected. During hydrothermal deactivation low rare-earth catalysts are more strongly dealuminated and the higher degree of dealumination results in a more pronounced mesopore volume, more non-framework alumina and the decreasing interaction between framework aluminium atoms amplifies the strengths of the active centres [19,20] and these catalyst properties enhance bottoms cracking. All these selectivity differences between the reaction severity variation by (i) catalyst mass and (ii) feed rate can be ascribed to the severity effects discussed above. Thus, this investigation highlights the decisive role of the catalyst-to-oil adjustment procedure in catalyst evaluations. Variation of catalyst-to-oil ratio by varying feed injection time has the advantage that the feed rate can be kept constant, and avoids some of ‘severity effects’ described above. However, this approach also has limitations which will be discussed in an upcoming paper. Briefly, variation by time-on-stream varies the level of coke-on-catalyst and thus the catalyst deactivation to obtain differences in reaction severity. To get a constant conversion comparison, a more active catalyst is compared at longer times-on-stream (longer deactivations) compared to a less active catalyst. Since catalyst deactivation also changes catalyst selectivity, this comparison has an effect on the selectivity measurements. In addition, in order to obtain ‘realistic’ catalyst-to-oil ratios, typical results are obtained at long times-on-stream (50–150 s) which tend to overemphasise the effects of matrix components on selectivity.

4. Comparison of results from different reactor types In order to determine the strengths and weaknesses of the latest micro-scale tests, this paper compares the

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D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

Table 4 Influence of different reaction severity adjustments on FCC catalyst rankinga Method C/O variation by

DCR Catalyst mass

SCT-MAT Catalyst mass

SCT-MAT Feed rate

ASTM-MAT Feed rate

C (*)

D (*)

C (*)

D (*)

C (*)

D (*)

Catalyst C (+)

D (+)

Activity at constant catalyst-to-oil ratio Conversion (%) 70

74

70

77

69

77

64

71

Yields at constant conversion of 70% (wt.% f f) Hydrogen 0.024 C1 +C2 1.9 Propylene 4.3 Propane 0.72 i-Butylene 2.3 i-Butane 2.3 C4 Olefinicityb 70.9 LPG 14.6 Gasoline 51.3 Light cycle oil 20.2 Heavy cycle oil 9.8 Coke 2.2

0.019 1.8 4.0 0.68 1.9 2.5 67.1 14.0 52.1 20.0 10.0 2.1

0.083 1.7 5.1 0.76 2.5 2.6 70.7 17.1 49.4 20.6 9.4 1.7

0.058 1.5 4.7 0.73 2.0 3.0 65.6 16.1 50.6 20.2 9.8 1.7

0.084 1.8 5.3 0.75 2.6 2.6 71.6 17.5 48.9 20.5 9.5 1.7

0.058 1.4 4.7 0.66 2.1 2.8 67.5 16.1 51.0 20.9 9.1 1.5

0.076 1.9 4.9 0.95 1.6 4.1 55.6 16.8 48.4 20.7 9.3 2.8

0.057 1.7 4.2 0.98 1.0 4.2 49.0 15.3 50.2 21.0 9.0 2.9

Gasoline composition (%) n-Paraffins i-Paraffins Aromatics Naphthenes Olefins

3.1 22.0 30.0 12.7 32.2

3.9 25.7 27.5 9.2 33.7

4.0 29.6 27.7 9.7 29.0

4.3 25.7 27.8 9.2 33.0

4.3 29.5 26.4 9.8 30.0

3.9 34.9 30.5 9.4 21.4

4.2 39.9 30.5 9.8 15.6

2.9 19.3 29.4 11.6 36.8

Catalyst properties

Catalyst C

Deactivation Al2 O3 (wt.%) RE2 O3 (wt.%) Zeolite surface area (m2 /g) Matrix surface area (m2 /g) Unit-Cell size (Å)

Hydrothermal: 5 h, 1089 K, 100% steam, 1 atm, fluidised bed 42.7 42.0 0.59 2.0 166 152 18 26 24.21 24.25

a b

Catalyst D

Feed properties: Table 2; curve fitting to activity and yield data was performed with (+) 4 observations, (*) 6 observations per catalyst. C4 composition in %.

results obtained with an improved fixed bed, and an improved FFB design. Sapre and Leib have shown that testing FCC catalysts which primarily differ in rare-earth content do not show the accurate ranking when tested in ASTM-type MAT’s and FFB units [5]. To re-evaluate these findings in the latest reactor designs, a set of commercially available FCC catalysts were examined in the SCT-MAT representing an annular diluted fixed bed and the FFB technique called advanced catalyst evaluation (ACE). For the validation of these techniques the performance of the same FCC catalysts was inves-

tigated in the Devison Circulating Riser (DCR) Pilot Unit. 4.1. Catalyst testing in the absence of metals Catalyst performance as a function of unit-cell size is well understood [21,22] and therefore such a comparison was used in this study. Three catalysts which only differed in the degree of rare-earth exchange were hydrothermally deactivated to unit-cell sizes typical for commercially equilibrated catalysts. These samples were tested in the DCR. Following

D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

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Table 5 Comparison of DCR, SCT-MAT and ACE testing at constant feed ratea,b Method reactor C/O variation by Feed pre-heat Catalyst time-on-stream

DCR (+) riser Catalyst mass Low (580–630 K) Constant (5 s)

SCT-MAT (*) fixed bed Catalyst mass Low (600 K) Constant (12 s)

ACE (+) fixed fluidised bed Time-on-streamc High (794 K) Variable (65–112 s)

Catalyst E

F

G

E

F

G

E

F

G

71

76

66

72

78

74

75

77

Yields at constant conversion of 75% (wt.% f f) Hydrogen 0.024 0.021 C1 +C2 1.8 1.7 Propylene 5.1 4.9 Propane 0.75 0.79 i-Butylene 2.2 1.9 i-Butane 3.2 3.3 C4 Olefinicityd 65.5 63.4 LPG 17.0 16.6 Gasoline 53.1 53.5 Light cycle oil 17.5 17.4 Heavy cycle oil 7.5 7.6 Coke 3.0 3.2 Coke-on-catalyst 0.30 0.41

0.015 1.5 4.4 0.77 1.5 3.4 60.3 15.4 54.6 16.9 8.1 3.3 0.58

0.083 1.7 5.4 0.87 2.4 3.2 66.9 18.4 52.9 15.1 9.9 1.9 0.60

0.063 1.6 5.2 0.86 2.0 3.6 63.1 18.0 53.4 15.0 10.0 2.0 0.71

0.051 1.6 5.0 0.97 1.5 3.9 57.8 17.4 53.9 14.5 10.5 2.1 0.90

0.080 1.7 5.4 1.1 1.9 5.2 52.0 19.2 51.1 18.2 6.8 2.9 0.55

0.065 1.8 5.0 1.2 1.4 5.7 45.5 18.7 51.4 18.3 6.7 3.0 0.70

0.042 2.1 4.8 1.8 1.1 6.7 38.6 20.1 48.7 16.5 8.5 4.1 1.14

Gasoline composition (%) n-Paraffins i-Paraffins Aromatics Naphthenes Olefins

3.0 22.4 31.1 11.4 32.1

3.2 23.8 32.3 11.9 28.9

3.5 26.4 32.5 12.7 24.8

4.1 28.8 28.5 8.5 30.0

4.2 32.0 29.3 8.5 25.9

4.6 34.3 29.5 9.0 22.6

3.0 35.3 30.0 7.7 24.0

3.0 38.2 32.3 7.6 18.8

2.9 40.0 28.8 7.8 14.6

Catalyst properties

Catalyst E

Catalyst F

Catalyst G

Deactivation RE2 O3 (wt.%) Al2 O3 (wt.%) Zeolite surface area (m2 /g) Matrix surface area (m2 /g) Unit-cell size (Å)

Hydrothermal: 5 h, 1089 K, 100% steam, 1 atm, fluidised bed 1.1 1.9 2.9 42.3 41.3 41.0 142 146 136 35 37 33 24.24 24.28 24.33

Activity at constant catalyst-to-oil ratio Conversion (%) 68

a

Feed properties: Table 2. Curve fitting to activity and yield data was performed with (+) 4, (*) 6 observations per catalyst. c Catalyst mass constant (9 g), variation of feed injection time at constant feed rate. d C composition in %. 4 b

the experiments, samples were withdrawn from the regenerator and investigated in a SCT-microactivity and an FFB unit. Key experimental settings and findings obtained by the three reactor types are summarised in Table 5. In each test protocol, catalyst activity was seen to increase with increasing unit-cell size as expected. Ranking agreement was also obtained for hydrogen, coke, product olefinicity, light

cycle-oil and heavy cycle-oil. However, the findings measured by SCT-microactivity testing were much closer to the DCR data than those obtained by the FFB (ACE) in terms of (i) absolute yields as well as (ii) relative yield differences between the catalysts. Much lower gasoline yields and product olefinicities were obtained by FFB testing than in the DCR and SCT-microactivity test. These deviations in FFB

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D. Wallenstein et al. / Applied Catalysis A: General 203 (2000) 23–36

testing from the DCR findings are ascribed to the very long catalyst time-on-stream in the FFB method leading to time averaging effects which are more pronounced than in the DCR and SCT-microactivity test. Moreover overcracking of the gasoline fraction occurs earlier in FFB’s than in plug flow reactors (fixed bed, riser reactors) owing to backmixing in the former [23,24], which also reduces gasoline selectivity and product olefinicity. The most striking discrepancy between the three reactor-types were the different rankings in catalyst performance observed for C1 –C4 compounds and gasoline selectivity. With increasing unit-cell size less C1 –C4 and more gasoline were produced by the DCR unit and the SCT-MAT as expected. The FFB technique, however, gave a reverse sequence for these fractions, contradicting well established knowledge. Formation of C1 –C4 compounds

and gasoline as a function of unit-cell size was shown by Pine [21] and Haas [22] and their findings have been verified by other research groups. Due to the relevance of this issue to the present work, some of the key findings obtained by previous investigations [21,22] are briefly reviewed. During hydrothermal deactivation low rare-earth catalysts are more strongly dealuminated than high rare-earths catalysts, hence, they have less active sites resulting in a lower activity. The higher degree of dealumination attenuates the interaction between framework aluminium atoms and therefore amplifies the acid strength of the active centres [19,20]. Protolytically induced reactions are enhanced by increasing acid strength, consequently more methane and ethane are formed at lower unit-cell sizes. Higher unit-cell sizes increase acid site density in the zeolite which increases

Table 6 Comparison of SCT-MAT and ACE testinga Method reactor C/O variation by Feed pre-heat Catalyst time-on-stream

ACE fixed fluidised bed Feed rate High (794 K) 100 s

SCT-MAT fixed bed Catalyst mass Low (600 K) 12 s

Catalyst H Activity at constant catalyst-to-oil ratio Conversion (%) Catalyst-to-oil ratio at 65% conversion Yields at constant conversion of 65% (wt.% f f) Hydrogen C1 +C2 LPG C4 Olefinicityb Gasoline LCO HCO Coke

I

H

I

65 5.9

68 4.8

69 1.8

78 1.3

0.62 1.9 12.3 64.1 41.4 16.9 18.1 8.8

0.96 1.9 11.8 64.1 40.5 18.8 16.2 9.8

0.31 1.4 14.1 70.2 45.1 12.5 22.5 4.1

0.30 1.3 12.7 78.9 46.7 13.4 21.6 3.9

Catalyst properties

Catalyst H

Catalyst I

Deactivation V (ppm) Ni (ppm) RE2 O3 (wt.%) Al2 O3 (wt.%) Unit-cell size (Å) Z-SA/M-SAc

Cyclic propylene steaming [17] 6500 6000 3400 3200 2.2 2.1 34.6 40.5 24.28 24.30 1.9 1.4

a

Feed properties: Table 2. Yield in %. c Zeolite surface area to matrix surface area ratio. b

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bimolecular hydrogen transfer reactions, and this leads to higher rates of conversion of gasoline range olefins to their analogous paraffins. Furthermore, since carbenium ions are terminated at higher molecular weight, gasoline yield is increased. Thus, the formation of C3 and C4 compounds decreases and gasoline yields increase at higher unit-cell sizes. The reverse ranking of the process described above was observed in FFB testing which poses a real problem for the evaluation of catalysts using this protocol. The conflicting catalyst performance in this latest FFB design may be ascribed to the substantial experimental discrepancies between DCR and FFB testing such as (i) feed pre-heat, (ii) adjustment of reaction severity, both discussed in Sections 3.2 and 3.3 and (iii) back mixing effects in the latter. The catalyst-to-oil ratio adjustment by variation of catalyst time-on-stream in the FFB unit results in averaging effects as a function of catalyst activity whereas this parameter is a constant in SCT-microactivity and the DCR testing. Moreover, very long catalyst times-on-stream are used in the FFB protocol which means that most of the feed is converted on a coke deactivated catalyst whereas in DCR and SCT-microactivity testing, the feed is contacted with a catalyst with decaying activity. The better agreement of catalyst performance in the fixed bed MAT and DCR can be rationalised by the fact

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that experimental settings are more alike in DCR and SCT-microactivity testing in terms of reaction temperature, feed preheat, catalyst time-on-stream and variation of reaction severity than in the FFB method. 4.2. Catalyst testing in the presence of metals The performance of two FCC catalysts with different matrix types and similar rare-earth-on-zeolite concentrations was investigated in the presence of metals. The catalysts were impregnated with vanadium and nickel by the incipient wetness method and deactivated by cyclic propylene steaming. As in the last section, the catalysts were compared both with the SCT-microactivity test, which does not have in-situ regeneration, and the FFB test (ACE), which does. The experimental settings and the findings are summarised in Table 6. Catalyst I is recommended for resid operation and therefore this catalyst is expected to convert more heavy cycle-oil than catalyst H and this was, indeed, observed in both protocols. The heavy cycle-oil to light cycle-oil ratio was lower for catalyst I than for catalyst H in both reactor types. The sequence for catalyst activity was also identical whereby the differentiation was more pronounced in the SCT-microactivity test. A reverse ranking in catalyst performance was observed for gasoline make.

Fig. 3. Coking rate of different catalysts as a dependence on catalyst contact-time.

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Catalyst I produced 2% (rel) more gasoline than catalyst H in the FFB and vice versa 3% (rel) less gasoline in the fixed bed. The ranking of FCC catalysts on the basis of gasoline selectivity in the SCT-MAT has been validated by comparisons with DCR data and therefore the FFB data are considered to be inaccurate. The inspection of coke yields showed the FFB unit to produce significantly more coke than the MAT (about 140% rel) and two reasons account for this finding: (i) the stronger catalyst decay due to the very long catalyst times-on-stream and (ii) the oxidation state of metals. To understand these results, experiments were performed in an ASTM MAT unit as a function of time-on-stream. The results (Fig. 3) show that the coke-on-catalyst concentration reaches its saturation point after 12 s time-on-stream. This implies that all subsequent reactions take place on a coke-deactivated catalyst. Consequently the time-averaged conversion at constant catalyst-to-oil ratio was lowest in FFB testing. Therefore a higher catalyst-to-oil ratio has to be used in the FFB test to achieve the conversions obtained by the SCT-microactivity unit and this was performed by reducing the feed mass at constant feed injection time. Fig. 3 shows that coke-on-catalyst is roughly constant after 12 s and therefore coke yields are primarily a function of the feed mass at longer catalyst times-on-stream. With decreasing feed mass, the coke yield increases and this partially accounts for the higher coke yields obtained with the FFB unit. The second contributor to the discrepancy in coke yields and one reason for the higher hydrogen make in the FFB unit are the different oxidation states of vanadium and nickel. In this case, the FFB used the in-situ regeneration technique described above to remove coke from the catalyst between the runs. This regeneration increases the oxidation state of the metals on the catalyst. Several authors have shown that the ratio of the dehydrogenation activity of nickel-to-vanadium is of the order 3–4 when these metals are in a reduced state. The catalytic activity of vanadium in the +5 state is 3–4 times higher than in lower oxidation states while the hydrogen and coke making properties of nickel are similar after deactivation in reducing or oxidising atmosphere [17]. Catalyst metals are oxidised in the FFB unit during the in-situ regeneration step, hence the dehydrogenation activity of vanadium is much higher in the second cracking experiment.

Fig. 4. (a) Influence of in-situ regeneration on the catalytic activity of vanadium and nickel: hydrogen selectivity following (i) in-situ regeneration (ACE) and (ii) replacing the catalyst (MAT) after each cracking experiment; (b) influence of in-situ regeneration on the catalytic activity of vanadium and nickel: coke selectivity following (i) in-situ regeneration (ACE) and (ii) replacing the catalyst (MAT) after each cracking experiment.

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In order to investigate this further, the same fresh catalyst was impregnated with vanadium and nickel to different levels. The non-metallated and the metallated samples were deactivated by cyclic propylene steaming and were tested in both FFB and ASTM-type MAT units. Hydrogen and coke selectivities are shown in Fig. 4a and b. A substantial increase in hydrogen and coke selectivities was observed for the metallated catalysts after the first in-situ regeneration step. This step change does not occur in MAT testing, since the coke loaded catalyst is replaced with fresh catalyst after each cracking experiment. The magnitude of the increase in coke and hydrogen yields after in-situ regeneration was more pronounced at high metal concentrations and no unexpected increase of coke and hydrogen make was observed for the metals-free catalysts. Thus, the enhanced hydrogen and coke make in the FFB unit is explained by the oxidation of vanadium during in-situ regeneration.

5. Conclusions The shortcomings of the conventional ASTM fixed bed method, which include channelling, coke profiles and poor isothermicity, have limited its effectiveness as a tool for predicting catalyst selectivity. Two divergent strategies have emerged to overcome these shortcomings: improved fixed bed testing (SCT-microactivity) and improved fluidised bed testing (FFB or ACE). In this paper, the strengths and weaknesses of these two approaches are explored. Experimental comparisons between the test units show that unit design and operating conditions have a profound influence on the outcome of catalyst assessments. The use of an innovative fixed bed design overcomes many of the limitations of the traditional ASTM reactor. Feed pre-heat also plays a decisive role in FCC and it has been shown that the simulation of the thermal shock that FCC feeds experience in riser reactors has a significant bearing on evaluating catalyst performance. Of utmost importance is the manner of adjusting reaction severity. The findings in this paper provide evidence that the method employed in the DCR Pilot Unit can be simulated only in microactivity testing by variation of catalyst mass. FFB based microactivity tests also provide good isothermicity, do not suffer from the development of

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a coke profile and channelling and certainly provide advantages with regard to feed dispersion. However, the decisive operating parameters for an accurate assessment of catalyst performance cannot be employed in these micro-versions of FCC units. The advantages of short catalyst times-on-stream cannot be utilised in FFB reactors since this would result in over-exaggerated fluidisation. On the other hand, longer contact times affect selectivities and can mask catalyst properties. Selectivities change significantly during catalyst time-on-stream resulting from intrinsic selectivity and shape selectivity changes caused by coke formation. In order to obtain improved efficiency of operation, reaction severity in FFB units has to be varied by feed rate or feed injection time. It has been shown that both techniques cannot match the performance in riser reactors. Moreover, the literature [4,5] describes plug-flow reactors as providing a better simulation of riser reactors than fixed fluidised beds due to (i) backmixing effects and (ii) the more pronounced time averaging in the latter. The findings obtained for metallated catalysts clearly show difficulties with the in-situ regeneration approach that can be used in either fixed or FFB reactors. Oxidation of vanadium during the first in-situ regeneration step dramatically enhances the catalytic activity and thus the dehydrogenation activity of vanadium. The improvements achieved by advanced metallation and deactivation techniques such as cyclic metallation with ageing in pilot units and cyclic propylene steaming cannot be utilised in the case of in-situ regeneration after the cracking experiments. Numerous comparisons of selectivity data obtained by (i) a conventional ASTM-type MAT, (ii) a SCT-MAT technique using an annular diluted fixed bed and variation of catalyst mass for adjustment of reaction severity and (iii) a riser pilot unit have been presented [9,10]. Excellent agreement of catalyst ranking and absolute yield levels with riser data were obtained by the new SCT approach and clearly, the only realistic catalyst evaluation is presently ensured by these two techniques. So far, such an agreement between riser and FFB units has not been described. This missing validation and the experimental findings and details discussed above question catalyst evaluations with the FFB technique.

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Acknowledgements The authors thank W.R. Grace & Co. for permission to publish this manuscript and T. Racke, H. Wolk, K. M. Smith and M. G. Hawkins for the experimental work.

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