Impact of rare earth concentration and matrix modification in FCC catalysts on their catalytic performance in a wide array of operational parameters

Applied Catalysis A: General 502 (2015) 27–41

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Impact of rare earth concentration and matrix modification in FCC catalysts on their catalytic performance in a wide array of operational parameters D. Wallenstein a,∗ , K. Schäfer a , R.H. Harding b a b

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

a r t i c l e

i n f o

Article history: Received 18 March 2015 Received in revised form 11 May 2015 Accepted 12 May 2015 Available online 21 May 2015 Keywords: FCC catalysts Unit cell size effects Matrix modification Gasoline formation Bottoms cracking

a b s t r a c t Fluid catalytic cracking (FCC) is the most flexible process in the refining industry to convert residues from the atmospheric distillation of crude oils into liquefied petroleum gas (LPG), gasoline and light cycle-oil (LCO). In order to meet the varying demands for these fractions via changes in the catalytic properties of FCC catalysts, the effects of varying rare earth content and matrix modification are primarily utilized. Rare earth content variation changes the behaviour of the zeolitic part in FCC catalysts with regard to its response to the hydrothermal deactivation and contaminant metals during FCCU operation; the rare earth content primarily determines the rate of dealumination and structural collapse of the zeolite and thus the resulting equilibrium unit cell size. This parameter strongly correlates with the rates of hydrogentransfer reactions, which in turn, have an impact on catalyst deactivation by coke formation. The degree of dealumination, structural collapse and coke-on-catalyst level affect the diffusion of the feed molecules into the zeolite and the residence time of the products in the zeolite. For the work presented here, FCC catalysts of different rare earth content were hydrothermally equilibrated in the absence and presence of contaminant metals. The interplay of the resulting structural properties with the catalytic performance was investigated by cracking a heavy vacuum gas-oil and a resid feed on these FCC catalysts at short catalyst time-on-stream and high temperature. With increasing rare earth content, the LCO selectivity decreased whereas gasoline and LPG selectivities run through maxima and minima respectively. The evaluation of the individual compounds in the gasoline fraction suggests that these observations could be attributed to mass transport limitations imposed by the structural changes of the zeolite as a function of rare earth content. However, the selectivity trade-offs which accompany the beneficial effects of increasing the rare earth levels on zeolite, as for example lower LCO selectivity, can be compensated by matrix modification. Thus all the advantages of high unit cell size catalysts such as high stability, high activity, low coke selectivity, low dry gas, high gasoline make of low olefinicity and gasoline sulfur reduction can be utilized. To put the findings into commercial perspective: amongst the criteria which impose constraints on FCCU operation are catalyst circulation, heat balance and from a legislative point of view the gasoline olefinicity. Therefore, comparisons of yields at constant catalyst-to-oil ratio, constant coke and constant product olefinicity are often more relevant to estimate commercial performance than the comparison of yields at constant conversion. Such an evaluation at constant catalyst-to-oil ratio shows substantial benefits in catalyst activity, bottoms cracking and gasoline yields at high rare-earth-on-FCC catalyst content whilst the desired reduction in gasoline olefinicity was achieved as well. © 2015 Published by Elsevier B.V.

1. Introduction Numerous studies reporting the effect of unit cell size on the product selectivity have been presented [1–9] but

∗ Corresponding author. Tel.: +49 6241 4031312; fax: +49 6241 403901344. E-mail address: [email protected] (D. Wallenstein). http://dx.doi.org/10.1016/j.apcata.2015.05.010 0926-860X/© 2015 Published by Elsevier B.V.

relatively few studies have been performed equilibrating the unit cell size and measuring the catalytic performance under FCCU operation-like conditions such as the presence of contaminant metals, catalyst deactivation in reduction–oxidation cycles, catalysts having age distribution, and gas-oil cracking at short catalyst times-on-stream and high temperatures. If such experimental parameters are not taken into account, the catalyst evaluations may be distorted with regard to their translation to

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D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

the performance in FCC units. Some examples are outlined below briefly:

In order to cover all these facets of catalyst deactivation and catalytic testing, an FCC catalyst series, where the catalysts only differed in rare earth content were metallated to high and moderate vanadium and nickel loadings to simulate commercial deactivation in resid and vacuum gas–oil operation. For the simulation of low metals operation non-metallated catalysts were used. The deactivation of both the metallated and metals-free samples was performed with reduction–oxidation cycles. Moreover, some of the data we had already published for catalysts having age distribution [16] and being relevant for the topic of this work were included. The work was accomplished by investigating the impact of matrix modification of the FCC catalyst series with the different rare earth contents on catalytic performance. Cracking experiments were performed in a microactivity test unit at short catalyst time-on-stream; the accuracy of this test had been validated by comparisons with data generated in riser pilot units [17]. The catalysts with the high metal loading were tested with a resid feed whilst for the testing of the moderately loaded and metals-free catalysts a heavy vacuum gas-oil was employed. For the interpretation of the catalytic performance, the pore volume distributions of the deactivated FCC catalysts and the complete product distribution of the C1 –C12 fraction were determined. The evaluation of the catalytic performance was carried out by considering several aspects. The first approach was the comparison of yields at constant conversion representing the classical one. However, there are many constraints in FCCU operation imposed by a wide array of mechanical limitations, downstream process capabilities and legislative product specifications. To give some examples, mechanical issues are wet-gas compressor capacity, catalyst circulation rate and heat balance, and therefore, changes in product distribution at constant conversion are not necessarily relevant for catalyst performance in commercial FCCU operation. In order to cover such aspects, it is also shown how data generated in laboratories can be discussed and interpreted for commercial applications.

• The presence of vanadium certainly plays an important role; it has been shown that the structural degradation of the catalyst occurs by the interplay of hydrothermal as well as by vanadium mediated dealumination of the zeolite and structural collapse. Moreover, other contaminant metals – for example nickel – accelerate catalyst deactivation during gas-oil cracking by coke formation [10–12]. • The physical properties of FCC catalyst particles equilibrated in FCC units vary from particle to particle as a function of their residence time in the FCC unit. Beyerlein et al. [13] and Palmer et al. [14] have shown that the catalytic performance of such equilibrium catalysts are dominated by the younger fractions which may change catalyst ranking in comparison to FCC catalysts without age distribution. • Changing the degree of dealumination via the variation of the deactivation severity rather than the variation of rare earth concentration may also affect mesoporosity and catalytic properties of FCC catalysts. We performed experiments to compare these two approaches and the findings revealed differences between the two procedures in terms of pore volume distribution and yield patterns obtained from microactivity testing; i.e. a consistent unit cell size reduction method has to be employed to investigate the impact of unit cell size on catalytic performance. • Catalytic performance testing at different catalyst times-onstream revealed that increasing times-on-stream de-emphasizes the contribution of zeolitic cracking and over-emphasizes the contribution of matrix cracking on product distribution [15]. The work reflecting best commercial conditions was performed by Pine et al. [1]; FCC catalysts with varying rare earth levels were hydrothermally deactivated without contaminant metals and a riser pilot unit was employed for selectivity testing. Moreover, the large sample size credits his work. However, the paper did not contain (i) any structural measurements of the deactivated catalysts, (ii) the distribution of the individual compounds in the gasoline range, and (iii) the LCO and coke selectivities; such pieces of information are necessary to rationalize the findings. There is, thus, a need to re-examine and re-evaluate the effects of varying the rare earth levels in several operation modes (low-, moderate- and highmetals operation, gas-oil and resid processing) and elucidating the differences in catalytic performance by a more detailed characterization of the physical catalyst properties and product distribution.

2. Experimental 2.1. Catalysts The FCC catalysts used for the experiments contain REUSY-type zeolites exchanged to different rare earth levels and dispersed in an alumina sol matrix; the properties are compiled in Table 1. These catalysts are of the same technology in view of matrix type as well as binder type. They are designed for low hydrogen and coke selectivities through nickel tolerance by use of a Ni trapping matrix.

Table 1 Catalyst properties. Catalyst Calcination Zeolite surface area, m2 /g Matrix surface area, m2 /g Unit cell size, Å Al2 O3 , wt% RE2 O3 , wt% Deactivation CPS without metals Zeolite surface area, m2 /g Matrix surface area, m2 /g Unit cell size, Å CPS with 3000 ppm V + 2000 ppm Ni Zeolite surface area, m2 /g Matrix surface area, m2 /g Unit cell size, Å CPS with 5400 pppm V + 2500 ppm Ni Zeolite surface area, m2 /g Matrix surface area, m2 /g Unit cell size, Å

A

B

C

D

E

F

G

241 72 24.51 47.3 1.1

246 62 24.53 47.0 1.8

227 68 24.56 49.1 2.2

226 68 24.58 48.3 3.3

245 60 24.59 47.1 3.7

252 53 24.63 45.0 5.7

200 51 24.64 55.1 4.9

141 39 24.25

147 35 24.29

137 30 24.32

149 33 24.35

152 33 24.36

159 24 24.44

124 21 24.46

131 43 24.24

145 33 24.27

134 37 24.31

142 30 24.34

144 32 24.36

152 30 24.41

121 26 24.46

126 35 24.23

131 38 24.28

137 36 24.33

146 28 24.36

145 21 24.42

112 18 24.46

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

Catalysts A–E were formulated with the same zeolite, matrix and binder content. The small differences in surface areas and Al2 O3 between these samples can be attributed to variations in catalyst preparation and in the analyses methods. Catalyst G was formulated with lower zeolite content than Catalysts A–E in order to compensate for its expected high activity. For balance kaolin was used. As a consequence Al2 O3 content of Catalyst G was higher and its zeolite surface area was lower. Since kaolin has very low cracking activity only a reduction of the overall cracking activity but no major impact on product selectivities reflecting the intrinsic catalytic performance of the zeolite were expected. Catalyst C was only available for the experiments with the non-metallated catalysts and the catalysts metallated to moderate vanadium and nickel loadings. 2.2. Catalyst pre-treatment The deactivation of the catalysts with metals was performed by cyclic propene steaming (CPS). The experimental conditions were as follows: The samples were calcined for 3 h at 540 ◦ C prior to loading with vanadium- and nickel-naphthenates using toluene as solvent. The quantity of the impregnation solution was double the amount of the FCC catalysts and the excess of solvent was removed in a rotary evaporator at 120 ◦ C under vacuum. The remaining organics were burned off with air at 250 ◦ C for 3 h followed by 700 ◦ C for 3 h in a shallow bed. Thereafter the samples were deactivated by CPS as described below. • Heat-up phase: room temperature to 716 ◦ C at 8.6 ◦ C/min, N2 purge. • Two reduction–oxidation cycles as described below during the heat-up phase from 716 ◦ C to 804 ◦ C at 1.1 ◦ C/min: 30 min; 50%, 5% propene in nitrogen + 50% steam 2 min; 50% nitrogen + 50% steam 6 min; 50%, 4000 ppm SO2 in air + 50% steam 2 min; 50% nitrogen + 50% steam • 29 reduction–oxidation cycles as described above at 804 ◦ C • 30 min; 50%, 5% propylene in nitrogen + 50% steam • Cool down under nitrogen flow. The coke-on-catalyst following CPS was burned off with air at 500 ◦ C, for 3 h in a shallow bed. Complete details and discussion of this deactivation protocol are given in Refs. [16,18–20]. The deactivation of the catalysts without metals was also performed by cyclic propene steaming (CPS) as described above with following modifications. The calcined samples were impregnated with a heavy vacuum gas-oil (HVGO) having high sulphur content; the characterization data are given in Table 2 (Feed A). This HVGO Table 2 Feedstock properties. Feed

A

B

C

Conradson carbon, wt% Average molecular weight Sulfur, wt% Caromatic , % Cparaffinic , % Cnaphthenic , % D-1160 simulated distillation IBP, ◦ C 5 vol%, ◦ C 20 vol%, ◦ C 40 vol%, ◦ C 60 vol%, ◦ C 80 vol%, ◦ C 95 vol%, ◦ C FBP, ◦ C

0.25 345 2.6 22.4 59.5 18.0

0.34 388 0.30 18.1 61.4 20.5

3.7 440 0.74 25.6 58.3 16.1

217 307 342 382 423 471 524 552

240 311 367 411 451 497 547 622

193 324 387 427 466 518 579 632

29

was diluted with toluene. The concentration of the HVGO in toluene (22%) was similar to that of the V/Ni naphthenates in toluene used for the metallation of the catalysts. The CPS procedure used in this work represents partial burn operation. Details are outlined in Refs. [16,19]; in these publications the durations of the reduction and oxidation steps are given. The ratios of the lengths of these oxidation-to-reduction steps classify the CPS method used as a partial burn protocol. FCCU operation in the presence of high contaminant metal (V, Ni) levels is predominantly performed in partial burn modes in order to reduce the deleterious effects of vanadium. Therefore catalyst deactivation is more governed by hydrothermal deactivation than by the effects of contaminant metals in partial burn operation. Hence this CPS method is also applied for testing of non-metallated catalysts. For these reasons this CPS procedure is also referred to a global deactivation method in previous publications.

2.3. Cracking experiments and analysis of products Cracking tests were performed with a heavy vacuum gasoil and an atmospheric resid feed, the properties are shown in Table 2. The experiments were performed with the Short Contact-Time Microactivity Test Unit (SCT-MATU) as described previously [17]. 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. In order to determine the product selectivities six cracking experiments at different reaction severities were conducted for each FCC catalyst to obtain conversions between 40 and 75%. The reaction severity was varied by changing the catalystto-oil ratio (C/O) where C/O is defined as the amount of catalyst divided by the total amount of feed. This ratio was varied by changing the mass of the catalyst while the amount of feed and the time-on-stream were kept constant at 1.5 g and 12 s respectively. Thermal effects and spatial changes in the catalyst bed were minimized by diluting of the FCC catalyst with glass beads of 0.2–0.3 mm diameter to a volume of 10 mL. The reaction temperature was 560 ◦ C. 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 560 ◦ C. The feed was injected into the catalyst bed and 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 mL/min nitrogen to a total gas volume (crack gas + nitrogen) of 600 mL which was determined by the amount of displaced water. After the cracking experiment the FCC catalyst was separated from the glass beads by sieving and the coke-on-catalyst was determined by a carbon analyzer type CS-344 [21]. The gas-chromatographic methods employed for analysis of the products obtained from the cracking experiments such as hydrogen, C1 –C6 gases, liquid compounds, gasoline composition, gasoline octane numbers have been described previously [21]. The yields were calculated as weight percent of reactant. The fractions of gasoline, light cycle-oil and heavy cycle-oil were determined by the cut points at 216 ◦ C and 338 ◦ C. Conversion is defined as:

100 wt% − (light cycle − oil, wt% + heavy cycle − oil, wt%).

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2.4. Data analysis To simplify the reporting the yields were compared at constant parameters such as conversion, coke and C/O ratio. The approach used for the evaluation of the cracking experiments and the computational method for the prediction of the yields at constant parameters has been described previously [22]. Briefly, conversions and yields were modelled with functions derived from reaction kinetics using C/O as the independent variable. Interpolated yields at constant conversion, coke, and C/O were calculated with the parameters obtained by regression analyses of the experimental data with these functions. The interpolated data represent yields at conversions in the steep part of the exponential function describing the dependence of conversion on C/O ratio [22]. Interpolations at higher conversions in the flat part of this curve could (i) misrepresent the corresponding C/O values, (ii) influence the ranking of selectivities due to overcracking effects and therefore the evaluation of the tested catalysts by the interpolated data may be distorted. 2.5. Physical characterization methods Surface area and pore volume distribution of the catalysts were determined by nitrogen sorption using a Micromeritics Tristar 3000TM unit. The zeolite and matrix surface areas were calculated by the t-plot method (Harkins/Jura) using the pressure range p/po = 0.06–0.35. The pore volume distribution was calculated from the desorption branch according to Barrett, Joyner and Halenda. The unit cell size of the zeolite Y was determined by X-ray diffraction (XRD) using a Bruker AXS D8 Advance analyzer according to the Standard ASTM D 3942-97 procedure. 3. Results and discussion 3.1. Effect of deactivation 3.1.1. Variation of (i) rare earth concentration at constant deactivation severity and (ii) deactivation severity at constant rare earth concentration Catalyst characterization data following CPS deactivation at constant deactivation severity were added to Table 1. For each catalyst series, the deactivation reduced the unit cell size in inverse relation to the amount of rare earth content in the catalysts. The zeolite surface areas generally increased with this parameter with the exception of Catalyst G. This catalyst had the highest rare earthon-zeolite concentration, and therefore, the highest unit cell size. In order to compensate for its high activity, this sample was formulated with lower zeolite content, i.e. the lower zeolite surface areas of Catalyst G following the different deactivation modes are attributable to the lower starting value. The pore volume distributions following the deactivation of metallated and metals-free catalysts are illustrated in Fig. 1a–c. The peak in the 100–400 A˚ range is attributed to the zeolite and the AUC (area under the curve) indicates a decrease in pore volume with increasing rare earth content and thus with increasing unit cell size. This finding is attributable to the stabilizing effect of rare earth compounds which led to lower dealumination rates and therefore, less mesopore formation occurred. Thus, the zeolite mesoporosity is less pronounced for high rare earth catalysts. The position of the zeolite peak showed a tendency to migrate to larger pore diameters with increasing rare earth level which was unexpected – the stabilization of the zeolite should diminish pore diameters. In order to investigate how the zeolite peak is shifted in case of reducing the unit cell size at constant rare earth content by variation of the deactivation severity, Catalyst E was CPS deactivated

Fig. 1. (a–c) Impact of unit cell size variation by different rare earth-on-catalyst levels on pore volume distribution in different deactivation modes.

at 10, 20, 40 and 80 h following metallation to the moderate level. The increase in deactivation time also led to a higher degree of dealumination indicated by unit cell size shrinkage and stronger structural collapse, both parameters and the corresponding zeolite surface areas are illustrated in Fig. 2. The increase in structural collapse at longer deactivation times is also mirrored by the increasing AUC’s. Moreover, the pore sizes were shifted to larger diameters with increasing dealumination rates which was not observed in the ‘rare earth approach’. Thus the different behaviour of the catalysts indicates that dealumination kinetics proceeded differently in the two dealumination routes. Increasing the rare earth

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

31

23

21

LCO [wt.% ff]

19

17

15

13

24.23 Å

24.28 Å

24.33 Å

24.36 Å

24.42 Å

24.46 Å

11

Fig. 2. Impact of unit cell size variation of catalyst E by different CPS deactivation durations on pore volume distribution; metal level: 3000 ppm V + 2000 ppm Ni.

content at constant deactivation severity lowered the structural collapse in terms of total mesoporosity indicated by the AUC’s in the pore volume distribution plots. However, the structural collapse appeared to be confined to certain regions where a stronger collapse occurred. One possible explanation is that the so-called self-healing of the zeolite by Si-reinsertion is delayed due to slower dealumination and structural collapse rates induced by the higher degree of rare earth stabilization. At constant deactivation time, fewer Si atoms are available for reinsertion and thus in regions of high Al-concentrations, the structural collapse is more accentuated leading to larger pores. However, the integral of the structural collapse is smaller at high rare earth content, indicated by the AUC’s. In the case of increasing the unit cell size at constant rare earthon-zeolite concentration by decreasing the time Catalyst E was exposed to deactivation, the reduction in structural collapse was indicated by both, a decrease in the AUC’s and a shift of the zeolite peak to smaller diameters. The decrease in deactivation time at constant rare earth stabilization, i.e. at a constant framework Al distribution and concentration in the zeolite, entails a lower degree of dealumination. Thus less structural collapse occurred and the mesoporosity is reduced with regard to both, pore size and total mesoporosity. The pore volume distribution patterns of the metals-free and metallated catalysts (Fig. 1a–c) were similar which is attributable to the deactivation mode. The CPS method employed represents a partial burn operation as already discussed in the Section 2. Therefore, the structural collapse of the zeolite by vanadium was less pronounced than would be expected in a deactivation protocol representing full burn mode. 3.2. Catalytic results 3.2.1. Variation of unit cell size by rare earth concentration; comparison of selectivities The unit cell size of the zeolite is an indicator of the degree of dealumination which is governed by the rare earth content and this parameter is used to describe and discuss the data in this section. The data obtained from microactivity testing following different catalyst pre-treatments and deactivations are compiled in Tables 3a–3c. As the unit cell size increased the activity was shifted to higher values. For product selectivities the following trends were noted: olefinicities in the C4 and gasoline fractions decreased as unit cell size was increased. Gasoline yields proceeded through ˚ The a maximum in the unit cell size range of about 24.33 A. inverse correlation was observed for the LPG fraction, which passed through a minimum at the same value. Bottoms cracking behaviour of the FCC catalysts was examined by LCO selectivities; the data in

35

45

55

65

75

85

Conversion [wt.% ff]

Fig. 3. Bottoms cracking represented by LCO formation as a function of unit cell size; deactivation: 5400 ppm V + 2500 ppm Ni, CPS; Feed: C.

Fig. 3 show the conversion of bottoms to decrease consistently with increasing unit cell size. In order to investigate the reason for the decrease in gasoline selectivity at higher unit cell sizes a detailed examination of the C-number distribution in the gasoline range was performed. For each C-number the compounds were grouped into n- and iparaffins, olefins, naphthenes and aromatics. The selectivity for olefins decreased consistently with increasing unit cell size at each C-number. The yields of n-paraffins and i-paraffins smaller than C8 generally increased with increasing the unit cell size whereas the larger compounds proceeded through maxima. In the unit cell size range of 24.24–24.33 A˚ an increase in C8 –C12 paraffins was generally observed whereas at higher unit cell sizes these yields decreased. The LPG formation as function of unit cell size showed a reversed trend strongly suggesting that the decrease in C8 –C12 paraffins at higher unit cell sizes is attributable to cracking reactions. Selectivities of naphthenes and aromatics >C8 generally display a similar trend; the cracking of side chains is considered to lower these selectivities. Thus the decrease in gasoline selectivity at high unit cell sizes can be ascribed to the cracking of olefins together with the larger paraffinic, naphthenic and aromatic compounds in the gasoline fraction. The trends in LCO and gasoline selectivities appear to contradict some recent literature where a constant increase in the gasoline selectivity and improvement of bottoms cracking with increasing unit cell size were reported [1,2]. However, those data were generated with a catalyst-to-oil adjustment by variation of the feed mass entailing artefacts by the so-called severity effects. Very detailed discussions of these artefacts and their impact on gas-oil cracking are given in Refs. [17,23]; severity effects lead to artefacts which do not allow the prediction of commercial catalyst performance by such data. Moreover, a re-evaluation of Pine’s work [1] which represents the classical reference of unit cell size effects on gasoline selectivity revealed that those data are consistent with our findings. The gasoline and LPG selectivities obtained by variation of unit cell size by rare earth content are given in Table 4 and this data also show a maximum for gasoline and a minimum for LPG as a function of unit cell size. Pine’s conclusion of a steady increase of gasoline selectivity with increasing unit cell size was based on unit cell size stabilization with different cations, whereas the data generated by variation of rare earth on FCC catalyst content alone confirmed our data. It is true that the maxima in gasoline selectivity as a function of unit cell size obtained in this work and in Pine’s work are

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D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

Table 3a Influence of rare earth-on-catalyst concentration on MAT selectivities Pre-treatment: 5400 ppm V + 2500 ppm Ni, CPS; MAT-testing with resid feed (Feed C). Catalyst

A

Unit cell size, Å 24.23 Rare earth, wt% 1.1 Yields interpolated at 65% conversion, wt% ff 0.47 H2 2.3 C1 + C2 C4 Olefinicity, % 80.6 LPG 12.4 44.5 Gasoline 20.3 LCO 14.7 HCO 5.0 Coke 1.5 Coke-on-catalyst, wt% Gasoline composition, wt% 4.5 n-Paraffins 19.4 i-Paraffins Aromatics 28.8 Naphthenes 9.1 Olefins 38.2 GC-MON 80.2 GC-RON 93.0 Activity 47 Conversion at c/o = 2.5 3.4 C/O at 65% Conversion C-Number distribution at 65% conversion, wt% ff Paraffins 0.87 n-C1 0.78 n-C2 n-C3 0.78 0.50 n-C4 n-C5 0.39 0.31 n-C6 0.22 n-C7 0.46 n-C8 n-C9 0.15 n-C10 0.16 n-C11 0.14 n-C12 0.19 1.01 i-C4 1.48 i-C5 i-C6 1.46 i-C7 1.04 i-C8 1.13 i-C9 1.03 i-C10 0.81 0.88 i-C11 i-C12 0.79 Olefins 0.69 C2 3.87 C3 4.13 n-C4 2.14 i-C4 34.1 % i-C4 in Tot.C4 3.21 n-C5 2.92 i-C5 47.7 % i-C5 in Tot.C5 4.72 C6 C7 3.92 1.52 C8 0.68 C9 Naphthenes 0.32 C6 0.66 C7 1.44 C8 1.13 C9 0.50 C10 Aromatics 0.22 C6 1.44 C7 3.03 C8 3.36 C9 2.24 C10 2.52 C11

B

D

E

F

G

24.28 1.8

24.33 3.3

24.36 3.7

24.42 5.7

24.46 4.9

0.44 2.1 76.2 12.1 45.5 19.1 15.9 4.9 1.7

0.32 1.8 72.8 11.7 47.2 18.3 16.7 4.6 2.4

0.27 1.7 71.9 12.3 46.3 17.2 17.8 4.7 2.7

0.24 1.7 68.6 12.8 45.7 16.3 18.7 4.6 2.7

0.19 1.7 62.7 13.1 45.4 16.1 18.9 4.7 2.8

4.8 21.8 28.7 9.3 35.5 80.0 92.4

4.8 25.7 27.7 9.5 32.4 79.7 91.6

4.9 25.9 27.5 9.8 31.8 79.7 91.5

4.9 27.4 27.2 9.5 31.0 79.8 91.4

4.8 28.5 27.9 9.5 29.2 79.7 91.0

59 2.9

74 1.9

76 1.8

76 1.7

77 1.7

0.80 0.68 0.71 0.54 0.42 0.35 0.24 0.47 0.16 0.17 0.15 0.20 1.29 1.58 1.66 1.19 1.28 1.23 0.94 0.98 1.04

0.70 0.59 0.59 0.52 0.44 0.36 0.25 0.48 0.19 0.19 0.15 0.19 1.57 2.11 2.32 1.47 1.41 1.42 0.98 1.04 1.36

0.68 0.53 0.60 0.55 0.46 0.38 0.28 0.46 0.19 0.19 0.14 0.17 1.71 2.34 2.46 1.62 1.45 1.36 0.94 0.92 0.91

0.68 0.52 0.67 0.63 0.50 0.41 0.26 0.43 0.19 0.17 0.13 0.15 1.99 2.72 2.73 1.66 1.58 1.36 0.86 0.89 0.74

0.67 0.52 0.67 0.70 0.50 0.42 0.28 0.42 0.19 0.15 0.13 0.11 2.48 2.89 2.93 1.83 1.74 1.31 0.82 0.84 0.58

0.62 3.69 4.02 1.83 31.3 3.15 2.67 45.9 4.51 3.78 1.44 0.60

0.54 3.42 3.94 1.65 29.5 3.18 2.46 43.6 4.45 3.28 1.34 0.55

0.51 3.66 4.16 1.61 27.9 3.15 2.45 43.8 4.17 3.18 1.23 0.55

0.53 3.79 4.18 1.54 26.9 3.07 2.38 43.7 4.13 2.97 1.10 0.53

0.53 3.94 3.92 1.43 26.7 2.87 2.19 43.2 3.76 2.84 1.11 0.48

0.34 0.76 1.45 1.14 0.52

0.43 0.88 1.43 1.29 0.47

0.46 1.06 1.41 1.16 0.46

0.46 1.03 1.43 1.00 0.42

0.49 1.16 1.35 0.94 0.39

0.23 1.44 3.04 3.43 2.31 2.59

0.25 1.37 3.16 3.63 2.19 2.44

0.24 1.49 3.15 3.46 2.21 2.17

0.27 1.44 3.10 3.33 2.11 2.19

0.27 1.59 3.36 3.36 2.09 2.01

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

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Table 3b Influence of rare earth-on-catalyst concentration on MAT selectivities pre-treatment: 3000 ppm V + 2000 ppm Ni, CPS; MAT-testing with HVGO (Feed B). Catalyst

A

Unit cell size, Å 24.24 Rare earth, wt% 1.1 Yields interpolated at 65% conversion, wt% ff 0.38 H2 1.5 C1 + C2 C4 Olefinicity, % 78.8 LPG 13.6 47.3 Gasoline 18.4 LCO 16.6 HCO 2.4 Coke 0.8 Coke-on-catalyst, wt% Gasoline composition, wt% 3.4 n-Paraffins 21.3 i-Paraffins Aromatics 27.4 Naphthenes 8.9 Olefins 38.9 GC-MON 80.8 GC-RON 93.8 Activity 53 Conversion at c/o = 2.5 3.0 C/O at 65% conversion C-Number distribution at 65% conversion, wt% ff Paraffins 0.57 n-C1 0.41 n-C2 n-C3 0.54 0.43 n-C4 n-C5 0.31 0.30 n-C6 0.17 n-C7 0.38 n-C8 n-C9 0.11 n-C10 0.10 n-C11 0.09 n-C12 0.15 1.42 i-C4 2.05 i-C5 i-C6 2.04 i-C7 1.26 i-C8 1.21 i-C9 1.10 i-C10 0.69 0.88 i-C11 i-C12 0.85 Olefins 0.53 C2 4.34 C3 4.68 n-C4 2.21 i-C4 32.1 % i-C4 in Tot.C4 3.50 n-C5 3.39 i-C5 49.2 % i-C5 in Tot.C5 5.24 C6 C7 4.01 1.51 C8 0.73 C9 Naphthenes 0.41 C6 0.95 C7 1.40 C8 1.07 C9 0.39 C10 Aromatics 0.18 C6 1.46 C7 3.27 C8 3.71 C9 1.89 C10 2.46 C11

B

C

D

E

F

G

24.27 1.8

24.31 2.2

24.34 3.3

24.36 3.7

24.41 5.7

24.46 4.9

0.36 1.4 74.9 13.3 47.8 17.6 17.4 2.4 1.0

0.28 1.3 72.7 12.7 48.6 17.4 17.6 2.3 1.0

0.27 1.2 70.5 12.8 48.6 17.0 18.0 2.3 1.2

0.22 1.2 68.4 13.3 47.9 17.0 18.0 2.1 1.2

0.15 1.2 67.0 13.8 47.6 16.0 19.0 2.0 1.4

0.16 1.1 62.4 14.4 46.9 15.5 19.5 2.4 1.3

3.7 24.0 27.6 9.8 35.0 80.2 92.5

4.0 26.7 27.7 9.7 32.0 80.3 92.2

4.1 28.5 27.8 9.7 30.0 80.2 91.7

4.2 28.7 26.9 10.6 29.6 80.0 91.3

4.5 30.1 26.2 10.3 28.9 80.0 91.3

4.4 33.0 25.0 10.8 26.8 80.5 91.3

65 2.4

69 2.2

73 1.9

76 1.8

80 1.4

76 1.9

0.52 0.38 0.49 0.45 0.35 0.32 0.17 0.42 0.12 0.12 0.10 0.15 1.76 2.18 2.22 1.50 1.46 1.22 0.82 0.91 1.13

0.48 0.34 0.47 0.47 0.39 0.34 0.20 0.47 0.14 0.13 0.11 0.15 1.87 2.51 2.73 1.55 1.63 1.37 0.90 1.02 1.25

0.45 0.32 0.44 0.47 0.41 0.36 0.22 0.46 0.15 0.13 0.11 0.14 2.05 2.67 2.98 1.80 1.92 1.54 1.05 0.99 0.92

0.42 0.32 0.49 0.54 0.44 0.35 0.25 0.47 0.15 0.12 0.10 0.12 2.26 2.99 2.93 1.80 1.92 1.39 1.01 0.96 0.77

0.40 0.31 0.46 0.56 0.49 0.41 0.29 0.48 0.15 0.12 0.10 0.09 2.43 3.28 3.36 1.97 1.73 1.40 0.96 0.95 0.68

0.33 0.28 0.49 0.65 0.55 0.41 0.26 0.43 0.15 0.11 0.09 0.07 2.90 4.04 3.88 2.26 1.56 1.34 0.88 0.90 0.59

0.49 4.04 4.56 2.03 30.8 3.30 2.99 47.5 4.57 3.77 1.41 0.66

0.44 3.64 4.43 1.80 28.9 3.22 2.87 47.1 4.34 3.26 1.25 0.61

0.43 3.77 4.36 1.66 27.6 3.08 2.67 46.4 3.99 3.00 1.25 0.58

0.45 3.94 4.43 1.62 26.8 2.99 2.55 46.6 3.85 3.02 1.22 0.55

0.45 4.26 4.45 1.64 26.9 3.02 2.55 45.7 3.77 2.77 1.16 0.50

0.48 4.49 4.34 1.54 26.2 2.88 2.38 45.2 3.43 2.46 0.95 0.49

0.49 1.09 1.44 1.25 0.40

0.49 1.16 1.49 1.13 0.45

0.56 1.29 1.39 1.04 0.42

0.60 1.35 1.51 1.16 0.44

0.61 1.44 1.44 1.05 0.38

0.77 1.45 1.44 1.04 0.34

0.17 1.43 3.39 3.73 2.00 2.48

0.17 1.42 3.32 3.90 2.11 2.54

0.18 1.43 3.37 3.86 2.10 2.58

0.17 1.39 3.24 3.67 2.05 2.37

0.17 1.43 3.02 3.49 2.02 2.34

0.20 1.68 2.70 2.99 2.03 2.15

34

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

Table 3c Influence of rare earth-on-catalyst concentration on MAT selectivities Pre-treatment: CPS without metals; MAT-testing with HVGO (Feed B). Catalyst

A

Unit cell size, Å 24.25 Rare earth, wt% 1.1 Yields interpolated at 65% conversion, wt% ff 0.067 H2 1.7 C1 + C2 C4 olefinicity, % 72.6 LPG 15.0 47.3 Gasoline 19.0 LCO 16.0 HCO 1.1 Coke 0.41 Coke-on-catalyst, wt% Gasoline composition, wt% 3.7 n-Paraffins 25.5 i-Paraffins Aromatics 29.1 Naphthenes 10.7 Olefins 31.0 GC-MON 80.8 GC-RON 92.7 Activity 62 Conversion at c/o = 2.5 2.7 C/O at 65% conversion C-Number distribution at 65% conversion, wt% ff Paraffins 0.57 n-C1 0.54 n-C2 n-C3 0.77 0.60 n-C4 n-C5 0.43 0.33 n-C6 0.20 n-C7 0.27 n-C8 n-C9 0.16 n-C10 0.13 n-C11 0.11 n-C12 0.11 2.04 i-C4 2.66 i-C5 i-C6 2.65 i-C7 1.24 i-C8 1.45 i-C9 1.34 i-C10 0.89 0.98 i-C11 i-C12 0.85 Olefins 0.59 C2 4.63 C3 4.87 n-C4 2.13 i-C4 30.4 % i-C4 in Tot.C4 3.13 n-C5 3.05 i-C5 49.4 % i-C5 in Tot.C5 4.07 C6 C7 2.73 1.22 C8 0.48 C9 Naphthenes 0.71 C6 1.09 C7 1.65 C8 1.21 C9 0.42 C10 Aromatics 0.10 C6 1.44 C7 3.64 C8 4.09 C9 1.98 C10 2.52 C11

B

C

D

E

F

G

24.29 1.8

24.32 2.2

24.35 3.3

24.36 3.7

24.44 5.7

24.46 4.9

0.052 1.5 68.3 14.3 47.7 18.6 16.4 0.9 0.42

0.044 1.3 64.6 14.1 49.4 18.5 16.5 1.1 0.58

0.040 1.3 64.6 13.8 49.7 18.2 16.8 0.9 0.56

0.036 1.2 63.3 14.2 49.9 18.2 16.8 0.9 0.56

0.029 1.4 62.2 14.4 48.5 18.2 16.8 1.2 0.80

0.028 1.4 57.5 15.8 45.5 17.6 17.4 1.2 1.0

3.9 28.2 28.5 10.4 29.0 80.5 92.0

4.3 30.4 29.7 10.1 25.6 80.8 91.4

4.4 31.0 30.0 10.0 24.5 80.3 91.1

4.6 32.3 30.1 9.6 23.4 80.1 90.4

4.5 32.0 31.8 9.3 22.4 80.6 90.8

4.5 32.5 31.8 9.3 21.9 80.7 90.7

70 2.1

73 1.9

74 1.6

75 1.6

75 1.5

77 1.2

0.51 0.49 0.68 0.63 0.47 0.37 0.21 0.30 0.16 0.14 0.11 0.12 2.34 2.91 3.32 1.40 1.59 1.39 0.95 1.00 0.86

0.45 0.40 0.72 0.68 0.52 0.42 0.23 0.36 0.18 0.14 0.11 0.15 2.61 3.29 3.78 1.62 1.81 1.43 0.98 1.07 1.01

0.43 0.39 0.67 0.67 0.51 0.43 0.24 0.39 0.19 0.15 0.12 0.15 2.57 3.28 3.61 1.76 2.05 1.52 1.09 1.11 1.01

0.41 0.36 0.69 0.70 0.54 0.43 0.27 0.45 0.20 0.14 0.13 0.15 2.72 3.54 3.81 1.78 2.23 1.57 1.08 1.11 0.99

0.45 0.40 0.74 0.63 0.54 0.42 0.25 0.39 0.19 0.14 0.12 0.12 2.83 3.67 3.72 1.77 2.07 1.51 0.95 1.00 0.83

0.50 0.44 0.90 0.91 0.58 0.43 0.24 0.36 0.18 0.14 0.11 0.08 3.32 4.16 3.77 1.63 1.92 1.38 0.82 0.97 0.79

0.53 4.23 4.6 1.80 28.1 3.03 2.81 48.1 3.85 2.51 1.16 0.46

0.48 4.05 4.45 1.55 25.8 2.91 2.75 48.6 3.40 2.16 1.00 0.41

0.47 4.01 4.41 1.49 25.3 2.86 2.65 48.1 3.31 2.03 0.95 0.40

0.46 4.17 4.49 1.41 23.9 2.78 2.57 48.0 3.07 2.00 0.84 0.42

0.50 4.51 4.43 1.27 22.3 2.70 2.42 47.3 2.86 1.80 0.68 0.40

0.49 4.96 4.49 1.24 21.6 2.72 2.34 46.3 2.58 1.72 0.66 0.36

0.75 1.15 1.55 1.13 0.37

0.80 1.30 1.48 1.05 0.34

0.80 1.29 1.44 1.08 0.35

0.74 1.36 1.34 1.04 0.29

0.72 1.31 1.30 0.92 0.27

0.67 1.31 1.29 0.88 0.27

0.08 1.37 3.48 3.98 2.02 2.67

0.08 1.56 4.05 4.02 2.17 2.78

0.08 1.43 3.86 4.41 2.27 2.88

0.08 1.48 3.86 4.15 2.27 3.14

0.08 1.51 4.13 4.77 2.14 2.81

0.13 1.56 4.07 4.55 2.07 2.70

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

Catalyst

A-1 A-2 A-3 A-4 A-5 A-6 A-7

RE2 O3 -0n catalyst, wt%

Unit cell size, A

0 0.2 0.6 0.9 1.0 1.3 2.5

24.23 24.24 24.25 24.26 24.25 24.28 24.34

Yields at 65% conversion Gasoline, wt%

LPG, wt%

58.2 60.0 60.3 61.2 59.1 60.3 60.2

13.3 11.0 10.7 10.6 12.9 11.2 11.0

different. This discrepancy can be attributed to the differences in catalyst technologies. The samples used in this work represent latest catalyst technologies and Pine’s work those from 1984 and earlier. Moreover the testing conditions (deactivation methods, catalytic testing) were quite different. The findings obtained in this work can be rationalized as follows: increasing the unit cell size had two effects; the number of active sites in the zeolite after dealumination is augmented and the mesoporosity in the zeolite decreased. These changes in zeolite properties reduced product olefinicities. The former depleted olefins by a higher rate of hydrogen-transfer reactions; the increased concentration of active sites accounts for this [2]. The latter increased diffusion limitations; therefore the residence time of the molecules in the zeolite increased which also favoured hydrogen-transfer reactions. Thus the incremental amount of olefins available for each incremental increase in unit cell size decreased significantly. Consequently the competition of olefins with paraffins at the active centres declined which is supposed to enhance paraffin cracking. The longer residence times of the molecules in high rare earth catalysts also supported paraffin cracking, which proceeds via energetically higher transition states than olefin cracking. The fact that only paraffins >C6 were converted can be explained by a nonclassical carbonium ion – carbenium ion reaction pathway, viz., protonated cyclopropyl (PCP) intermediates. Briefly, for isomerization reactions of n-alkanes in HF-SbF5 and on Pt/CaY zeolite catalysts PCP’s were postulated [24,25]. This mechanism circumvents the formation of primary carbenium ions which would be necessary in the case of n-alkane isomerization via the classical hybrid and methyl-shifts. Moreover, PCP’s are hybrids of resonance structures and therefore, the protonation of alkanes proceeding via this mechanism have lower energetic maxima than those via the ‘classical’ pentacoordinated carbonium ions. S. Tiong Sie translated the proposed existence of PCP’s in isomerization reactions into the FCC process, and published two excellent pieces of work with regard to the role of PCP’s in FCC cracking [26,27]. N-hexane is the smallest molecule which can be cracked via PCP’s. That the reactivity of C8 > C7 > C6 towards cracking reactions increased much sharper than proportional to the carbon number is explained by carbenium ion chemistry – the concerted action of PCP ring opening and ␤-scission allows the formation of a secondary carbenium ion in case of n-octane cracking, a primary one for n-heptane and a methylcarbenium ion for n-hexane. Cracking reactions via PCP’s supply a linear olefin and a branched paraffin and the product distribution obtained in the C4 and C5 olefin range give further evidence for the PCP mechanism. It is shown in Tables 3a–3c that C4 and C5 i-olefins were shifted away from their thermodynamic equilibria (i-C4 : 45%, i-C5 : 55% [7]) with increasing unit cell size; vice versa the unbranched olefins were shifted above their equilibrium concentrations (n-C4 : 55%, n-C5 : 45% [7]). Thus the PCP mechanism theory is supported by the experimental data observed in this work too. The findings suggest that the gasoline fraction was stabilized ˚ up to this by hydrogen-transfer up to a unit cell size of 24.33 A;

value sufficient olefins were available to occupy the active centres and to be converted by hydrogen-transfer to more crack-refractory compounds of the same C-number. Above this value the low concentration of olefins and the longer residence times of molecules in the zeolite enhanced paraffin cracking to such a rate that gasoline yields declined. The decrease in bottoms cracking with increasing unit cell size can be interpreted by several hypotheses; the changes in (i) zeolite mesoporosity, (ii) non framework alumina, (iii) strength of the acid centres, and (iv) hydrogen-transfer in the HCO (heavy cycleoil) fraction. The decrease in mesoporosity with increasing unit cell size reduced the externally available surface area of the zeolite for the conversion of large HCO molecules, and this reduced their cracking rate. The lower strength of the acid centres and the lower amount of non-framework alumina at high unit cell sizes have been reported to impair bottoms cracking [28,29]. The increase in hydrogen-transfer might convert naphthenic and paraffinic rings attached to polyaromatics to aromatic rings lowering the crackability of the HCO fraction. The latter hypothesis is supported in Fig. 4 which displays LCO selectivity as a function of C4 -olefinicity. C4 -olefinicity is the most robust proxy for hydrogen-transfer in fluid catalytic cracking and Fig. 4 shows that LCO formation and hydrogen-transfer activity are strongly coupled with each other. This strong correlation suggests that the effects of unit cell size, strength of the acid centres, non framework alumina, mesoporosity, and mass transport limitations on bottoms cracking can be described by one parameter, the C4 -olefinicity. This concept is supported by the literature which has shown that there are more parameters than unit cell size which affect hydrogen-transfer. Increase of hydrogen-transfer at constant unit cell size as a function of mesoporosity has been shown by Corma et al. [30]. In their work the performance of hydrothermally and chemically (SiCl4 ) dealuminated zeolites were compared. The latter did not show a partial structural collapse due to the reinsertion of Si in the vacancies which are formed during dealumination. Therefore the mesoporosity in the chemically dealuminated zeolite was far less pronounced which prolonged the residence times of the molecules in the zeolite. Consequently, depletion of olefins by hydrogen-transfer and paraffin cracking at constant unit cell size was enhanced in the case of chemically dealuminated zeolites. The role of mass transport limitations in hydrocarbon cracking was also reported in [29,31,32] and these papers support the conclusions of Corma’s work. Both the results described in the literature cited above and our findings suggest that hydrogen-transfer is governed by several parameters and that hydrogen-transfer describing indices can be used as a unified indicator to classify the bottoms cracking characteristic of FCC catalysts.

21

LCO at 65% Conversion [wt% ff]

Table 4 Data published by PINE et al. [1].

35

20

19

18

17

16

3000ppm V + 2000ppm Ni, CPS; Feed B (HVGO) 5400ppm V + 2500ppm Ni, CPS; Feed C (RF)

15 60

65

70

75

80

C4-Olefinicity at 65% Conversion [%] Fig. 4. LCO as a function of hydrogen-transfer activity.

85

36

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

Table 5 Metal levels and deactivation durations for preparation of FCC catalysts with age distribution. Fraction no.

1

2

3

4

5

6

Average

V, ppm Ni, ppm CPS, h

500 150 5

1000 400 10

2000 1000 20

3000 1800 40

4000 2800 70

5000 4000 110

2583 1692 42.5

Coke-on-catalyst formation following the different pretreatments and cracking tests with different feeds was quite different and increased in the following sequence: High metal loading, resid feed > moderate metal loading, HVGO feed > metals-free deactivation, HVGO feed. Notwithstanding, the trends in catalytic performance observed as a function of rare earth content were rather similar; i.e. no mass transport restrictions with visible consequences were imposed on the catalyst by coke formation. There are several reasons which could account for this. The formation of coke is classified in 4 categories [12]: (i) Contaminant coke (coke formed by catalytically active metals such as V, Ni, Fe, Cu), (ii) Carbon residue, this coke is directly related to the conradson carbon in the feed, (iii) Catalytic coke results from the reactions of the feed on the zeolite and from secondary reactions (hydrogen-transfer), (iv) Cat-to-oil coke is formed from the unstripped hydrocarbons. The findings suggest that not much of the contaminant coke and carbon residue were formed during the bulk of catalytic cracking reactions at the active centres of the zeolite, the zeolite openings or inside the channels. Most of these two coke-types appear to be formed at locations were they did not hinder the diffusion of the feed molecules and that of the products into, inside and out of the zeolite. Another possible explanation provides the formation of the cat-to-oil coke. This coke-type is mainly formed from the compounds which are difficult to remove from the catalyst in the stripping step. Hence the cat-to-oil coke is mainly formed after the bulk mass of feed is already converted. Vanadium and nickel on catalyst and increasing feed heaviness impede the stripping of hydrocarbons but primarily after the bulk mass of feed is converted, i.e. in the subsequent stripping step. Thus the cracking reactions appeared mainly to be governed by the zeolitic part of the catalyst not being hindered by coke formation. Therefore, it can be concluded that the differences in the yield patterns can be ascribed primarily to the variation in unit cell size and the structural changes of the FCC catalyst which accompany unit cell size variation. The catalyst-to-oil ratios given in Tables 3a–3c show that the low rare earth catalysts require much more catalyst to achieve similar conversion levels as the high rare-earth catalysts. One could argue that the higher LCO formation at low unit cell size can be attributed to the higher catalyst mass in the reactor compared to the high unit cell size catalysts. Since however secondary reactions such as hydrogen-transfer increase with increasing unit cell size (although the catalyst-to-oil ratio decreases, i.e. less catalyst is used to achieve the same conversion) it is illogical to argue that LCO formation would decrease due to the decreasing catalyst mass in the reactor. Therefore, the LCO and hydrogen-transfer shifts seen in this work are to be attributed intrinsic catalyst properties and not to any artefacts introduced by varying catalyst mass in the reactor to achieve similar conversion levels. 3.2.2. Age distribution FCC catalysts equilibrated in FCC units contain continuums of catalyst and metal ages whereas deactivation with the CPS method

produce uniformly metallated and deactivated catalysts. Hence it is of interest to investigate whether there are consequences of this difference on catalyst performance. For this purpose catalysts deactivated in a CPS method producing catalysts with age distribution were compared with catalysts deactivated with a conventional CPS method. These data had already been published [16] and only the data relevant for the topic of this work are given. The catalysts with age distribution were prepared with the metallation and CPS method described in the experimental section. For each catalyst six age fractions were prepared by impregnating the catalysts to six different vanadium and nickel levels followed by CPS deactivations of different durations; the corresponding vanadium and nickel levels and deactivation durations are shown in Table 5. The six age fractions of the corresponding catalysts were blended following deactivation to equal parts (16.6% from each fraction). For comparison catalysts without age distribution were metallated to the same metal levels and deactivated with the CPS method described above. Physical properties and the key results for catalytic performance are compiled in Table 6. Unit cell size as a function of CPS deactivation duration and metal levels are shown in Fig. 5a and the pore

Fig. 5. (a–b) FCC catalysts deactivated with age distribution: Unit cell sizes and pore volume distribution of the individual age fractions.

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

37

Table 6 Influence of rare earth-on-catalyst concentration on MAT selectivities Pre-treatment: CPS deactivated catalysts with and without age-distribution MAT-testing with HVGO (Feed B). Catalyst

H

Calcination 275 Zeolite SA, m2 /g Matrix SA, m2 /g 57 Unit cell size, Å 24.53 Al2 O3 , wt% 50.4 RE2 O3 , wt% 1.1 FCC catalysts with age distribution (see Table 5) Unit cell size, Å 24.26 141 Zeolite-SA 38 Matrix-SA Yields interpolated at 65% conversion, wt% ff 0.17 H2 1.7 C1 + C2 C4 olefinicity, % 66.2 13.9 LPG 46.7 Gasoline LCO 19.1 15.9 HCO Coke 2.0 0.71 Coke-on-catalyst, wt% Activity 60 Conversion at c/o = 2.5 2.8 C/O at 65% conversion FCC catalysts without age distribution Unit cell size, Å 24.26 Zeolite-SA 160 39 Matrix-SA Yields interpolated at 65% conversion, wt% ff 0.27 H2 1.8 C1 + C2 72.1 C4 Olefinicity, % 13.9 LPG Gasoline 46.6 19.0 LCO 16.0 HCO 2.1 Coke 0.83 Coke-on-catalyst, wt% Activity 64 Conversion at c/o = 2.5 C/O at 65% conversion 2.5

I

J

K

L

261 57 24.56 46.4 2.0

237 41 24.61 47.3 2.8

246 60 24.58 46.9 3.7

252 53 24.63 45.0 5.9

24.28 133 33

24.31 110 30

24.34 124 32

24.44 104 29

0.15 1.5 61.7 13.8 47.2 19.0 16.0 2.1 0.84

0.14 1.4 60.0 13.1 48.1 18.7 16.3 2.1 1.1

0.15 1.5 58.2 13.0 48.3 18.3 16.7 2.3 1.3

0.12 1.5 54.1 13.2 47.3 17.6 17.4 2.5 1.4

65 2.5

71 1.9

72 1.8

71 1.8

24.28 153 41

24.32 138 32

24.36 147 33

24.46 115 26

0.28 1.6 65.6 13.7 47.0 18.8 16.2 2.1 1.1

0.20 1.5 61.3 13.6 47.7 18.1 16.9 2.0 1.2

0.18 1.3 60.2 13.2 48.0 17.6 17.4 1.6 1.2

0.14 1.5 56.1 13.8 46.9 17.0 18.1 1.9 1.4

72 2.0

74 1.7

75 1.3

76 1.3

volume distributions of the individual age fractions for one catalyst in Fig. 5b. These data demonstrate the laboratory deactivated catalysts with age distribution to contain continuums of catalyst ages like those of FCC catalysts equilibrated in FCC units. The catalytic results obtained for the samples with and without age distribution were evaluated from two perspectives; first in terms of differences in the levels of activity and selectivities between the two deactivation modes and second the impact of rare earth variation on activity and selectivity. The conversions obtained at constant catalyst-to-oil ratio (Fig. 6a) shows the activities of the samples with age distribution to be lower, which can be attributed to their lower zeolite surface areas vs. those of the samples without age distribution. The higher hydrogen-transfer rates of the samples with age distribution (C4 olefinicity, Table 6) over the samples without age distribution can be ascribed to the catalytic dominance of the younger fractions in these samples. The higher gasoline selectivity (Fig. 6b) of the samples with age distribution is a consequence of a better gasoline stabilization due to their higher hydrogen-transfer activity. Bottoms cracking activity is reflected by the LCO selectivity in Fig. 6c. This parameter is more pronounced for the samples with age distribution, which is related to the high bottoms cracking potential of the older fractions in these catalysts. Regarding the impact of rare earth content in FCC catalyst variation on catalyst performance the following trends were

seen for both, the samples with and without age distribution. With increasing rare earth content the overall cracking activity and hydrogen-transfer rates increased, gasoline selectivity went through a maximum and bottoms cracking declined. Hence the described effects of unit cell size on the catalytic performance of FCC catalysts apply to a wide array of deactivation conditions and feedstocks.

3.2.3. Variation of unit cell size by variation of deactivation severity; comparison of selectivities It is shown in section 3.1.1 that the variation of unit cell size by (i) variation in deactivation severity at constant rare earth content and (ii) variation of rare earth content at constant deactivation conditions led to different structural properties. MAT-selectivities obtained from two catalysts out of each series are compiled in Table 7. For each series the expected ranking in terms of activity LPG, gasoline, product olefinicity and LCO was obtained, however this ranking is distorted when the two series are merged to one series. Activity, hydrogen, product olefinicity and LCO produced by the catalysts deactivated at 30 h CPS duration did not fit into the ranking of the 20 h series and the differences in textural properties could account for this. Thus this data suggests that unit cell size effects on product distribution can be distorted if different strategies for unit cell size variation are employed.

38

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

a) Catalyst Activity (Conversion) Conversion at Constant C/O [wt.% ff]

80

76

72

68

64

60 with Age Distribution

Single Point Deactivation

56 0

1

2

3

4

5

6

7

Rare-Earth on FCC Catalyst Content [wt.%]

b) Gasoline

Fig. 7. Impact of matrix modification on pore volume distribution; Deactivation: CPS with 5400 ppm V + 2500 ppm Ni.

a) LCO-formation as a function of unit cell size Deactivation: 3000ppm V + 2000ppm Ni, CPS, MAT testing: Feed: B

48.5

48.0

19

47.5

47.0

46.5 with Age Distribution

Single Point Deactivation

46.0 0

1

2

3

4

5

6

7

Rare-Earth on FCC Catalyst Content [wt.%]

LCO at 65% Conversion [wt.% ff]

Gasoline at 65% Conversion [wt.% ff]

49.0

18

17

16 Base 15 24.20

24.25

24.30

24.35

24.40

24.45

19.5

Unit Cell Size [Å]

19.0

b) LCO-formation as a function of hydrogen-transfer activity Deactivation: 3000ppm V + 2000ppm Ni, CPS, MAT testing: Feed: B 19

18.5

18.0

17.5

17.0 with Age Distribution

Single Point Deactivation

16.5 0

1

2

3

4

5

6

7

LCO at 65% Conversion [wt.% ff]

LCO at 65% Conversion [wt.% ff]

c) LCO

Additional Alumina Matrix

18

17

16

Rare-Earth on FCC Catalyst Content [wt.%]

Base

Additional Alumina Matrix

15

Fig. 6. (a–c) Impact of age distribution of FCC catalysts of different rare earth contents on conversion, gasoline and LCO selectivities.

65

67

69

71

73

75

77

79

C4-Olefinicity at 65% Conversion [%]

c) LCO-formation as a function of hydrogen-transfer activity Deactivation: 5400ppm V + 2500ppm Ni, CPS, MAT testing: Feed: C 21

LCO at 65% Conversion [wt.% ff]

3.2.4. Matrix modification An increase of the zeolite unit cell size reduces mesoporosity and non-framework alumina. The former imposes mass transport limitations on the feed and products in the zeolite, the latter reduces cracking activity. The data in the preceding sections suggest that both can reduce bottoms cracking. In this section it is shown how this trade-off can be compensated by matrix technology. The incorporation of special alumina matrices in FCC catalysts has been shown to improve bottoms cracking [33], the pre-cracking of the feed by the matrix improves the accessibility of the molecules to the interior of the zeolite and the additional alumina could work like non-framework alumina. Moreover, the pre-cracking of the feed by the matrix generates a more olefinic feed [34]; i.e. feed of a higher crackability enters the zeolite. In order to investigate this issue, some of the catalysts employed in the experiments in the preceding

20

19

18

17

16 Base

Additional Alumina Matrix

15 66

68

70

72

74

76

78

80

82

C4-Olefinicity at 65% Conversion [%] Fig. 8. (a–c) Impact of matrix modification on LCO-formation.

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

39

Table 7 Influence of unit cell size adjusted by (i) rare earth-on-catalyst concentration and (ii) variation of deactivation time on MAT selectivities. Pre-treatment: CPS without metals; MAT-testing with HVGO (Feed B). Catalyst

A

A

B

B

Unit cell size, Å Zeolite-SA CPS duration Data interpolated at: 65% conversion H2 C1 + C2 C4 Olefinicity LPG Gasoline LCO HCO Coke Coke-on-catalyst Gasoline composition n-Paraffins i-Paraffins Aromatics Naphthenes Olefins Activity Conversion at: c/o = 3 C/O (g/g) at: 65% conv.

m2 /g

24.24 130 30 h

24.25 141 20 h

24.28 131 30 h

24.29 147 20 h

wt%ff wt%ff % wt%ff wt%ff wt%ff wt%ff wt%ff wt%ff

0.066 1.7 70.4 15.4 46.8 18.7 16.3 1.1 0.39

0.067 1.7 72.7 15.1 47.1 19.0 16.0 1.1 0.35

0.053 1.6 66.4 15.1 47.3 18.0 17.0 1.0 0.41

0.052 1.5 68.3 14.3 48.2 18.6 16.4 0.9 0.38

% % % % %

3.9 25.5 29.5 10.4 30.7

3.9 25.6 29.7 10.4 30.4

4.1 29.0 29.7 9.8 27.5

4.0 27.8 28.5 10.6 29.1

%

69 2.7

65 3.0

73 2.4

71 2.6

sections were modified by a special alumina matrix. The matrix component was formulated into the catalyst – the bottoms cracking functionality of this matrix has been shown for both, lab testing and commercial operation [33]. Catalyst deactivation and testing was carried out as for the catalysts reported in the preceding sections. Pore volume distributions of (i) the catalysts formulated without additional alumina matrix (Catalysts E and H) and (ii) the analogous matrix containing species are illustrated in Fig. 7. This graph shows this matrix to generate a significant amount of additional pores in the 60–150 A˚ range. The bottoms cracking obtained by such matrix modified catalyst-types and the analogous ‘matrix-free’ counterparts is demonstrated by the LCO selectivities as a function of unit cell size in Fig. 8a. This data exhibits a substantial improvement in bottoms cracking by the alumina matrix over the whole unit cell size range. Apparently the matrix modified the feed prior entering the zeolite in terms of (i) reducing the average size of feed molecules and (ii) enhances its olefinicity and both are conducive to bottoms

cracking. Thus these data demonstrate clearly that the conversion of bottoms can be enhanced easily by matrix modification. 3.2.5. Consideration of FCCU operation constraints and legislative product specifications for catalyst evaluation MAT results can also be analyzed on a constant coke yield basis which better represents the heat balanced constrained FCCU than the comparison of selectivities. Favourable coke selectivity usually allows the refinery to operate the FCCU at higher conversions, provided catalyst circulation rate is not an additional limitation (in such a case a comparison at constant C/O ratio will also provide useful information). An increase in coke selectivity at higher rare earth content was reported in Refs. [3,35]. In this work the coke selectivity did not increase up to a unit cell size of 24.42 A˚ in the presence of V and Ni and the manufacturing route of the FCC catalysts employed may account for this improvement in performance. The high rare earth catalysts used in these experiments

Table 8 Influence of rare earth-on-catalyst concentration on MAT yield distribution at constant coke and constant C/O. Pre-treatment: 3000 ppm V + 2000 ppm Ni, CPS; MAT-testing with HVGO (Feed B). Catalyst

C

Unit cell size, Å Rare earth, wt% Data interpolated at: H2 C1 + C2 C4 Olefinicity LPG Gasoline LCO HCO Coke Coke-on-catalyst Gasoline composition n-Paraffins i-Paraffins Aromatics Naphthenes Olefins Activity Conversion C/O (g/g)

E

F

C

E

F

24.41 5.7

24.41 5.7

0.15 1.2 73.6 13.9 48.8 15.9 18.0 2.0 1.4

24.31 2.2 Constant c/o: 2.5 0.31 1.4 69.9 13.8 50.6 16.7 14.6 2.6 1.1

24.36 3.7

wt%ff wt%ff % wt%ff wt%ff wt%ff wt%ff wt%ff wt%ff

24.31 24.36 2.2 3.7 Constant coke: 2 wt% ff 0.26 0.22 1.1 1.2 79.6 74.8 11.5 13.3 46.1 48.2 18.0 17.0 21.1 18.1 2.0 2.0 1.0 1.1

0.32 1.8 58.8 17.2 53.3 14.5 9.4 3.5 1.4

0.30 2.1 50.1 19.3 54.4 12.5 7.2 4.2 1.7

% % % % %

3.9 24.3 26.8 10.2 34.7

4.1 28.5 26.8 10.7 29.8

4.2 29.8 26.9 10.4 28.7

4.1 27.5 29.0 9.4 30.0

4.5 33.1 31.9 9.3 21.3

4.8 37.1 34.6 8.4 15.1

% 2.0

61 1.8

65 1.4

66 2.5

69 2.5

76 2.5

80

40

D. Wallenstein et al. / Applied Catalysis A: General 502 (2015) 27–41

contain a novel faujasite zeolite, produced by a process giving the zeolite excellent hydrothermal stability, metals tolerance and thus, low coke selectivity [36]. That the catalysts reported in this work were (i) tested in the presence of contaminant metals and (ii) deactivated with reduction–oxidation cycles may also favourably contribute to the coke and stability issue [16,18,19]; the data reported in the literature and showing the poor coke selectivity at high rare earth content were measured on hydrothermally deactivated catalysts without metals. Yields interpolated at constant coke and constant C/O ratio are compiled in Table 8. Both comparisons show that the increase of the unit cell size from 24.31 to 24.41 A˚ dramatically reduced bottom yields and yielded more gasoline and LPG. The compositional changes in the gasoline fraction – despite the different conversions – were directionally similar to the changes observed at constant conversion. The decrease in gasoline olefinicity was even more accentuated. Thus the novel high rare earth catalysts (with improved coke selectivity over the catalysts reported in [3,35]) provide the refinery a catalytic route to comply with legislative gasoline specifications [37,38], which require reductions in olefin content, without imposing penalties in other valuable products. Another interesting aspect regarding gasoline olefins reduction provides the comparison of bottoms cracking by FCC catalysts formulated without and with an additional alumina matrix at constant olefinicity – such data are visualized in Fig. 8b and c. These graphs show that the lower the product olefinicity the larger the distance in LCO yields between the ‘matrix-free’ and matrix containing catalysts; i.e. the beneficial impact of matrix modification on bottoms cracking is significantly higher in case of zeolites with a high hydrogen-transfer activity. From the previous sections it is clear that increasing the unit cell size; or in other words, the increase in hydrogen-transfer activity decreases bottoms cracking at constant conversion. The results described in this section imply that there appears to be more room for bottoms cracking improvement by matrix modification at a high hydrogen-transfer activity, i.e. at a low product olefinicitiy. Here, the impact of (i) pre-cracking the feed by the matrix and (ii) providing a more olefinic feed to the zeolite on bottoms cracking was more pronounced than at lower unit cell sizes where diffusion limitations play a minor role and more olefins are available.

4. Conclusions The differences in rare earth-on-zeolite indirectly exert their influence on catalytic performance by changing the number of acid centres, the mesoporosity and non-framework alumina content. The decrease in LCO selectivities with increasing unit cell size and the maxima in gasoline selectivity in the range of 24.32–24.36 A˚ lend strong support to our perception that mass transport limitations is the overriding criterion in catalytic cracking. The mass transport in the zeolite allowed the gasoline to increase up to ˚ above this point mass transport an unit cell size of ∼24.33 A, limitations enhanced the rate of paraffin cracking and thus the gasoline selectivity was lowered. The selectivities of the larger LCO molecules were governed by mass transport limitations over the whole unit cell size range but it is shown that this trade-off can be adjusted easily by matrix modification. It is also shown that hydrogen-transfer reflecting indices (e.g. C4 -olefinicity, gasoline olefinicity) reflect the interplay of the textural changes with catalytic performance very well and that such indices can be used to explain differences in gasoline and LCO formation; e.g. the beneficial impact of matrix modification on bottoms cracking. The different metal-loadings and feeds employed in this body of work cover resid and vacuum gas-oil operation. The textural

properties of the catalysts obtained by lab simulation of the different FCCU deactivation modes were rather similar and this is attributable to the partial burn mode of the regeneration step in the CPS method used in this work; the structural degradation of the zeolite is more governed by hydrothermal deactivation rather than by vanadium even in the resid mode. Coke-on-catalyst levels were quite different and increased towards resid operation. Surprisingly the yield patterns and catalyst ranking in the three modes were quite similar; i.e. the higher coke-on-catalyst levels did not lead to any visible effects by mass transport limitations. The findings suggest that most of the additional coke, formed as a consequence of the shift from vacuum gas-oil towards resid operation, appears not to be formed in the zeolitic part of the FCC catalyst but on the matrix. Another possibility is the formation of coke from the unstripped hydrocarbons following the bulk of the catalytic cracking and hydrogen-transfer reactions whereby resid feed and metals increase this kind of coke formation. Moreover, the short contact-times used for the experiments may also account for the unchanged yield pattern, such results would certainly not be obtained by testing in fixed fluidized beds (FFB). The FFB-MAT technique over-emphasizes the cracking activity of matrices due to the extremely long catalyst times-on-stream [23]. Another item was the approach of unit cell size variation for a systematic investigation of unit cell size effects on catalytic performance. It is shown that the adjustment of unit cell size by two different approaches, (i) the variation in deactivation severity at constant rare earth content and (ii) variation of rare earth content at constant deactivation conditions for one catalyst series to be tested can distort the evaluation. Finally, several aspects with regard to the translation of labtesting results into commercial operation were shown by examples. The classical approach of comparing the product selectivities alone does not reflect commercial needs where FCCU hardware and legislative requirements impose restrictions on FCCU operation and the product pattern. Comparisons of a product fraction at constant coke, catalyst-to-oil ratio or C4 -olefinicity are more appropriate for such evaluations. The comparisons at these parameters demonstrated the tremendous advantages of high unit cell size catalysts much clearer than yields at constant conversion, assuming that zeolite and the finished FCC catalysts were manufactured by appropriate processes. Thus these FCCU-operation related evaluations accentuate more realistic the benefits in stability (activity retention), coke, dry gas, gasoline yields, gasoline olefinicity, gasoline sulfur content and bottoms conversion which can be gained by high unit cell size catalysts than the classical selectivity approach. Acknowledgement The authors thank W.R. Grace & Co for permission to publish this manuscript. References [1] L.A. Pine, P.J. Maher, W.A. Wachter, J. Catal. 85 (1984) 466–476. [2] A. Haas, J.R.D. Nee, Erdöl, Erdgas, Kohle 112 (1996) 312–3146. [3] G.W. Young, W. Suarez, T.G. Roberie, C. Wu-Cheng, Reformulated gasoline: the role of current and future FCC catalysts, in: Presented at the NPRA Meeting, March 17–19, San Antonio, TX, 1991. [4] K. Rajagopalan, A.W. Peters, J. Catal. 106 (1987) 410–416. [5] D.M. Stockwell, The role of porosity in the cracking selectivity of FCC catalysts, in: International Conference on Refining Processing, AIChE Conference Proceedings, Atlanta, 5th–9th March, 2000, pp. 132–136. [6] S.J. Miller, C.R. Hsieh, Octane enhancement in catalytic cracking by using highsilica zeolites, American Chemical Society Symposium Ser, 1991, pp. 452. [7] C. Wu-Cheng, W. Suarez, G.W. Young, The Effect of Catalyst Properties on the Selectivities of Isobutene and Isoamylene in FCC, vol. 88, AIChE Symp. Series, No. 291, 1992. [8] C. Wu-Cheng, E.T. Habib, K. Rajagopalan, T.G. Roberie, R.F. Wormsbecher, M.S. Ziebarth, in: G. Ertl, H. Knözinger, F. Schüth, J. Weitkamp (Eds.), Fluid Catalytic

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