Comparison between fixed fluidized bed (FFB) and batch fluidized bed reactors in the evaluation of FCC catalysts

Chemical Engineering Journal 183 (2012) 433–447

Contents lists available at SciVerse ScienceDirect

Chemical Engineering Journal journal homepage: www.elsevier.com/locate/cej

Comparison between fixed fluidized bed (FFB) and batch fluidized bed reactors in the evaluation of FCC catalysts Francisco Passamonti a , Gabriela de la Puente a , William Gilbert b , Edisson Morgado b , Ulises Sedran a,∗ a b

Instituto de Investigaciones en Catálisis y Petroquímica INCAPE (FIQ, UNL–CONICET), Santiago del Estero 2654, S3000AOJ Santa Fe, Argentina PETROBRAS S.A./CENPES, R & D Center, Av. Horacio Macedo, 950, Cidade Universitaria, Ilha do Fundão, 21941-915 Rio de Janeiro, RJ, Brazil

a r t i c l e

i n f o

Article history: Received 26 July 2011 Received in revised form 10 November 2011 Accepted 20 December 2011 Keywords: FCC Catalyst evaluation FFB CREC Riser Simulator DCR pilot plant

a b s t r a c t The laboratory reactors ACE fixed fluidized bed FFB and batch fluidized bed CREC Riser Simulator were compared in the conversion of two commercial vacuum gas oil feedstocks (paraffinic and aromatic types) over two equilibrium commercial FCC catalysts (octane-barrel and resid types) under similar conditions. Reaction temperatures were 510 and 540 ◦ C and the catalyst to oil mass ratios were from 4 to 9 in the FFB unit, with time on stream varying from 112 to 50 s, and 6.35 in the CREC Riser Simulator, with reaction times ranging from 5 to 25 s. The results were contrasted in some cases to those of a DCR circulating pilot plant unit at 540 ◦ C. When the different product yields were considered, the same catalyst ranking was observed in both laboratory reactors on the whole, differences between catalysts and particularly between feedstocks being more perceptible in the FFB reactor. The CREC Riser Simulator reactor showed a linear yield curve for gasoline, which facilitated the analysis and comparisons; in this case, an overcracking regime was not shown, like in the case of the FFB reactor. The LCO yield curves were more defined and differences between catalysts developed more clearly in the FFB reactor. Coke yields were very high in the FFB reactor, typical of confined beds with continuous feed of the reactants, while in the CREC Riser Simulator reactor they were in the range of those observed in the pilot plant and commercial units. The two laboratory reactors showed complementing potential in the laboratory evaluation of commercial catalysts and feedstocks for the FCC process. © 2011 Elsevier B.V. All rights reserved.

1. Introduction FCC is the major conversion process in many refineries, converting the VGO range (350–550 ◦ C) portion of the crude oil, which represents about one third of the total refinery capacity, mostly to high value C3–C10 hydrocarbons which constitute the majority of the motor gasoline fuel produced and a variety of intermediate feedstocks for the petrochemical industry and for other high grade fuel production processes (e.g. isomerization, alkylation and MTBE synthesis) [1]. The need of continuous replacement of old by new catalysts particles, in order to maintain activity, provides an opportunity for substantial change in the FCC yield profile by the relatively simple change of the catalyst system, without having to wait for the end of the unit turnaround. This may be used by the refiner to adjust to changes in the fuel market or to changes in feedstock quality in a short time frame, thereby capturing differences in prices and

∗ Corresponding author. E-mail address: usedran@fiq.unl.edu.ar (U. Sedran). 1385-8947/$ – see front matter © 2011 Elsevier B.V. All rights reserved. doi:10.1016/j.cej.2011.12.081

improving the overall business profitability. Catalyst manufacturers have been introducing new technologies over the years and special emphasis is given to the fine tuning of the catalyst systems according to the characteristics of the particular FCC unit. The strong impact of the catalyst on the global performance of the FCC unit and its profitability justifies the effort to guarantee the use of the best formulation available, and creates a demand for proper catalyst testing methodologies. Then, the procedure for the selection and evaluation of the FCC catalysts is critical [2]. Moreover, catalyst and process developments also call for a suitable laboratory tool to help in an evaluation as close to reality as possible. The high complexity and extreme magnitude of the commercial process severely complicate its faithful reproduction in the laboratory [3]. It is widely accepted that pilot circulating riser units reproduce more closely the specific environment of FCC commercial units [4], but they require both high investment and operating costs and their operation is quite complicated [5,6]. DCR Davison Circulating Riser units [7] are an example of these setups. Thus, laboratory tests are still the most commonly used methods to characterize the performance of FCC catalysts. Most of the laboratory tests are performed on MicroActivity Test (MAT, ASTM D-3907/03)-type fixed bed reactors, for which neat

434

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

advantages (e.g., ease of construction and operation) and disadvantages (e.g., operation mode and reactant–catalyst contact) are apparent. Undoubtedly, MAT-type reactors have been the standard in FCC-related laboratories [8], with a large number of different configurations and operative approaches being used [9–11]. A similar test methodology can be applied to flow reactors where a fluidized bed is confined (FFB [12]). In this case the bed nature is closer to that of commercial units, but the catalyst particles, which deactivate continuously by coke deposition, similarly to the case of MAT fixed bed reactors, see always a non changing, fresh feedstock. FFB reactors have become very familiar in FCC laboratories and tend to be the new standard setup [13]. The CREC Riser Simulator laboratory reactor [14] is the basis for the construction of an alternative approach for FCC catalyst and feedstock evaluation, as well as for some issues in process development. The reactor, with a fluidized bed of catalyst, ideally mimics the riser reactor in commercial units, following the analogy between position in the riser and reaction or contact time in the laboratory unit. It has been used extensively in FCC-related research, such as the modelling of kinetics, diffusion, and adsorption [15,16], the testing of new operative modes [17–19] and the assessment of particular product yields [20]. The performance of MAT units has been compared to that of pilot plant units, showing that for the same catalysts and feedstocks, many differences can be established in observed catalyst rankings and selectivities, depending on the catalyst type and deactivation procedures; most important differences were that, at constant conversion, a DCR pilot plant unit produced more olefins (both in gasoline and light gas) than the MAT unit, and that gasoline yields were lower and coke and LCO yields were higher in the MAT for active matrix catalysts [7]. Wallenstein et al. [21] demonstrated that by modifying the MAT technique it is possible to eliminate the ranking reversals identified in the comparison of catalyst selectivities observed in the ASTM-MAT and riser pilot units. The comparison between the performances of a MAT reactor and the CREC Riser Simulator reactor in the evaluation of two commercial VGOs using three equilibrium catalysts showed that product yield structures were very different [22]. The yield curves of the main hydrocarbon groups as a function of conversion followed linear behaviors in the fluidized bed reactor, while a non-linear dependence on conversion was observed for the yields in the MAT reactor, particularly in the cases of gasoline and coke. Moreover, the consequences on product quality derived from the use of so different devices and contact modes were that, for example, the naphtha obtained in the CREC Riser Simulator reactor was more paraffinic and less aromatic than the one obtained with the MAT reactor; ranks of catalysts based on the various hydrocarbon fractions observed in the naphtha from each setup also differed in most of the cases [23]. It is the objective of this manuscript to compare the results obtained in the conversion of two commercial vacuum gas oil feedstocks over two equilibrium commercial FCC catalysts under similar conditions in laboratory FFB and CREC Riser Simulator reactors, in terms of conversions and various product yields. In order to validate overall results, the observations in the laboratory reactors were also compared to those from a DCR Davison circulating pilot plant unit.

2. Materials and methods Two equilibrium commercial FCC catalysts were used, their properties being shown in Table 1. For use in the CREC Riser Simulator, catalyst particles were sieved and the fraction larger than 100 ␮m was used; however, all the properties in Table 1 were essentially the same for both fractions.

Table 1 Properties of the catalysts. Property

Units

Catalyst

X-ray fluorescence composition SiO2 Al2 O3 RE2 O3 P2 O5 Na2 O Fe2 O3 TiO2 SO4 Ni V Cu Sb Physical properties Apparent bulk density Y zeolite crystallinity (XRD) Unit cell size, Ao (XRD) Accessibility (AAI)a BET SA Micropore SA (t-plot) External SA (t-plot) Particle size distribution <149 ␮m <105 ␮m <80 ␮m <40 ␮m <20 ␮m

% % % % % % % % ppm ppm ppm ppm

E-cat R

E-cat L

50.8 43.4 2.43 1.46 0.53 0.66 0.32 0.11 1204 894 28 46

60.7 34.9 0.63 1.58 0.47 0.81 0.30 0.13 3520 174 9 19

g/mL % nm a.u. m2 /g m2 /g m2 /g

0.88 26 2.426 15.0 178 120 59

% % % % %

93 70 45 3 0

0.81 27 2.425 4.5 153 126 27 90 61 32 1 0

a The AAI (Akzo Accessibility Index) was assessed following the method by Hakuli et al. [47].

The properties of the feedstocks used, which were two commercial vacuum gas oils with different characteristics, are shown in Table 2; VGO-L (paraffinic) has high API density and concentration of saturated compounds, while VGO-R (aromatic) has low API density and high concentration of aromatic compounds. The laboratory reactors used were the CREC Riser Simulator [14] and an ACE (Model R+, supplied by M/s Kayser Technology Inc., USA) fixed fluidized bed unit (FFB) [24].

Table 2 Properties of the feedstocks. Property

Feedstock VGO-L Paraffinic

Density 20/4 (g/cm3 ) API (◦ ) Distillation (◦ C) 0 v% 10 v% 30 v% 50 v% 70 v% 90 v% Final Total sulfur (%) Basic nitrogen (ppm) Aniline point (◦ C) Ramsbottom carbon residue, RCR (%) Saturates (%) Monoaromatics (%) Diaromatics (%) Triaromatics (%) Polyaromatics (%) Viscosity (cStk) 60 ◦ C 82.2 ◦ C 100 ◦ C

0.8984 25.3 223.1 369.8 431.3 477.4 537.3 701.8 750.0 0.215 578 107.6 1.71 60.9 14.2 16.0 5.7 3.1 37.7 16.7 2.0

VGO-R Aromatic 0.9328 19.6 317.4 392.4 439.6 469.6 502.4 543.0 597.0 0.534 1014 83.6 0.43 47.4 18.3 21.0 8.6 4.8 73.9 27.2 14.8

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

435

Fig. 1. CREC Riser Simulator reactor: (a) basic design concept, (b) schematic representation.

The CREC Riser Simulator (RS) reactor has a turbine on top of a chamber that holds the catalyst bed between porous metal plates. The turbine rotates at 7500 rpm, thus inducing a low pressure area in the upper central zone in the reactor that makes gases recirculate in the upwards direction through the chamber, thus fluidizing the catalyst bed. When the reactor is at the desired experimental conditions, the reactant is fed hot at 80 ◦ C with a 0.25 ␮L syringe through an injection port and vaporizes instantly, thus setting the initial time. After the desired reaction time is reached, the gaseous mixture is evacuated immediately and products can be sent to analysis. Additional descriptive details can be found in e.g. [25]. A schematic representation of the unit is shown in Fig. 1. In the automated FFB ACE unit, the feed is sprayed inside a 9 g confined fluidized bed of catalyst at a rate of 1.2 g/min, which is kept at the stabilized reaction temperature by an electric heater, together with a stream of 100 mL/min of nitrogen. The reactor has two concentric tubes, catalyst being confined as a fluidized bed in the annular space between the tubes. The feed is injected through the internal tube that functions also as a preheater. At the injection point the feed meets a nitrogen upward flow that helps in the catalyst fluidization and prevents the feed from impinging on the reactor bottom. After reaction, the catalyst is stripped by nitrogen at 140 mL/min during additional 360–740 s (the higher the injection time the higher the stripping time) to remove adsorbed reaction products. The product stream leaving the reactor is partially condensed and the liquid and gas products are collected and analyzed separately. A schematic representation of the FFB reactor is shown in Fig. 2. Reaction temperatures were 510 and 540 ◦ C in both laboratory units, and all the catalysts were pre-calcined at 600 ◦ C during 1 h

in air before testing. Catalyst to oil mass ratio (C/O) in the RS unit was 6.35, achieved with a catalyst mass of 0.8 g and a reactants mass of 0.126 g, and reaction times were 5, 10, 15, 20 and 25 s. As usual in flow reactors, C/O ratio in the FFB unit was cumulative: it is the relationship between catalyst mass and the total mass of reactants fed in the experiment (mass flow rate times time on stream); thus the C/Os were 4, 5.5, 7 and 9, corresponding to times on stream of 112, 82, 64 and 50 s, respectively. In all the cases the reaction products were analyzed by gas chromatography. Mass balances (recoveries) in the RS reactor closed to more than about 94% in all the cases, while in the FFB unit they were in the 101–105% range. Coke yields were assessed by a thermal programmed combustion and methanation method in the catalyst samples from the RS reactor, and by in situ coke burning with air at 695 ◦ C for 25 s, followed by CO2 quantification in the FFB. Some experiments were also performed in a DCR Davison Circulating Riser pilot plant unit. The DCR unit is a circulating pilot riser with a 2 m long adiabatic riser. The C/O relationship, defined as the ratio between the catalyst circulation and feed rates, was controlled by fixing the riser outlet and regenerator temperatures, and changing the feed pre-heat temperature; different pre-heat temperatures were used for each catalyst to generate product yield versus conversion curves for interpolation and comparison at constant coke or constant conversion between different catalysts. Riser and regenerator temperatures were 540 and 730 ◦ C, respectively. Feed rate was 1000 g/h, reactor pressure was 25 psig, and the feed temperature was changed from 120 to 370 ◦ C. Mass balances in the unit were calculated from the liquid product weight, gas product volume and GC observed composition, and coke mass calculated from the flue

436

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

Fig. 2. Fixed fluidized bed reactor. Schematic representation: (1) catalyst hopper, (2) catalyst load line, (3) feed tank, (4) feed pump, (5) reactor, (6) feed line, (7) electric furnace, (8) CO converter (coke burn), (9) liquid condenser and collector, (10) gas collector, (11) gas GC analysis, (12) gas volume measurement by displaced water, (13) nitrogen/air inlet.

gas volume and CO2 , CO and O2 concentrations. Mass balances in the DCR were within 100–103 wt%. In all the cases dry gas, LPG, gasoline, LCO and bottoms yields were calculated from the GC analysis or simulated distillation, using 221 ◦ C and 434 ◦ C as cutpoints for gasoline and LCO, respectively. All data points were analyzed by multiple linear regression analysis using the software STATISTICA 8.0. The forward stepwise method was applied for the confidence level of 95%, without considering interactions between dependent variables. 3. Results and discussion The different contact modes between reactants and catalyst beds in the laboratory reactors, as well as the different flow models, impose very different actual meanings for the experimental parameters time and C/O. In effect, in the RS reactor, time is contact, residence or reaction time, while in the FFB unit it refers to time on stream. Moreover, the C/O relationship in the RS unit is the actual ratio between catalyst and reactant masses, whereas in the FFB unit it is usually assessed as depicted in Section 2. The range of reaction times in each unit was aimed at achieving a set of comparable conversions. Considering that the objective of this paper is to discuss similarities and differences in the performances of different laboratory setups for given feedstocks and equilibrium catalysts under similar conditions, and since a strong uncertainty rules the simulation of the process of equilibration of fresh catalyst particles to obtain equilibrium catalysts in the laboratory [3,26], equilibrium catalyst samples taken from refineries were used (see Table 1). They represent a conventional catalyst aimed at maximizing gasoline octane-barrels (E-cat L) and a catalyst formulated to process heavier feedstocks and to minimize bottoms yields (E-cat R). With this aim, the latter catalyst possesses a higher intake of active matrix, as inferred from its external surface area. Their close micropore

area and XRD crystallinity suggest similar contents of Y-zeolite in the equilibrium state. Although they contain quite different rare earth content, the unit cell size of the remaining Y-zeolite, associated to the acid site density, became approximately the same. In both catalysts, an additive to boost propylene yield was added in a proportion corresponding to nearly 1 wt.% ZSM-5 in the fresh state. 3.1. Conversion Conversion can be defined in different ways in FCC: typically, it is the sum of the yields of dry gas (DG), LPG, gasoline, LCO and coke [27], and sometimes the yield of LCO is excluded. In this manuscript, conversion was defined as x (%) =



yi = yDG + yLPG + yGasol + yLCO + yCoke

i

where yi refers to the yields (wt.%) of the different hydrocarbon groups. Moreover, it was confirmed that no significant differences can be observed in the conclusions in case of using the other usual definition. The conversions observed for the two feedstocks over the two catalysts, in each of the laboratory reactors at both temperatures, are shown in Fig. 3. As expected in the RS reactor, which is a batch, closed reactor, conversion increased steadily as a function of reaction time, with a decreasing rate. In this reactor, the bed is fluidized by the own reacting mixture, and the catalyst particles meet the same portion of reactants while reactions take place – initially the VGO, and then, during the elapsing reaction time, the mixture of products and unconverted feedstock. This would be the ideal contact in a plug flow riser unit. Conversion profiles in the FFB reactor were linearly decreasing as a function of time (time on stream). In the FFB reactor, the same amount of fluidized catalyst particles, which are initially clean, and deactivate continuously by coke deposition during the experiment,

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

100

100

(a)

(b) 90

Conversion. %wt

Conversion. %wt

90 80 70

80 70 60

60 50

0

5

10

15

20

25

50

30

40

60

Time, s

100

120

100

120

100

(c)

(d)

90

Conversion. %wt

90

80 70

80 70

60 50

80

Time, s

100

Conversion. %wt

437

60

0

5

10

15

20

25

50

30

Time, s

40

60

80

Time, s

100

(e) Conversion. %wt

90 80 70 60 50 2

3

4

5

6

7

8

CTO Fig. 3. Conversion profiles as a function of time in RS and FFB reactors and conversions in DCR tests as a function of C/O ratio: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols: () E-cat R/VGO-R, () E-cat R/VGO-L, () E-cat L/VGO-R, () E-cat L/VGO-L.

are always contacted by fresh feed. Then, conversion should drop as a function of time. However, in contrast to MAT-type reactors, activity profiles do not develop inside the reactor but the bed behaves homogeneously. The conversions reached in each of the laboratory reactors are somewhat above commercial values, but they could be adjusted properly by simply modifying some experimental parameters. The different reaction times in the RS reactor opened to a more extended range of conversions. It can be seen in both reactors that for a given catalyst the feedstock VGO-L is more reactive than the feedstock VGO-R at all conditions (see Fig. 3); this is consistent with the paraffinic character of VGO-L and the aromatic character of VGO-R. The comparison between catalysts, for the same feedstock and conditions in both reactors, shows that catalyst E-cat R is more active than catalyst

E-cat L; however, at the highest temperature of 540 ◦ C, differences between catalysts are less noticeable if the paraffinic VGO-L is used, and for the case of the RS reactor they are hardly distinguishable. The impact of reaction temperature was observed to be more significant in the RS reactor, with the exception of the pair E-cat R–VGO-L. 3.2. Product yields The yields of dry gas are shown in Fig. 4. It can be seen that the yield curves of the group suggest that it is a stable primary plus secondary product, as deduced from the results in the RS reactor, which covered a more extended conversion range. Results in the FFB reactor, which correspond to the high conversion zone, have an important slope as a function of conversion, very similar to the one observed in the RS reactor. While in the FFB reactor the yields

438

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

5.5

5.5

(a)

4.5

4.5

4.0

4.0

3.5 3.0 2.5 2.0

3.0 2.5 2.0 1.5

1.0

1.0 60

70

80

90

100

Conversion, %wt 5.0

0.5 50

70

80

90

100

90

100

5.5

(c)

(d)

5.0 4.5

4.0

4.0

Dry Gas, %wt

4.5

3.5 3.0 2.5 2.0

3.5 3.0 2.5 2.0

1.5

1.5

1.0

1.0

0.5 50

60

Conversion, %wt

5.5

Dry Gas, %wt

3.5

1.5 0.5 50

(b)

5.0

Dry Gas, %wt

Dry Gas. %wt

5.0

60

70

80

90

100

0.5 50

60

Conversion, %wt

70

80

Conversion, %wt

5.5 5.0 4.5

(e)

Dry Gas. %wt

4.0 3.5 3.0 2.5 2.0 1.5 1.0 0.5 50

60

70

80

90

100

Conversion, %wt Fig. 4. Yield curves for dry gas: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols as in Fig. 3.

were different according to the feedstock – VGO-R yielding more dry gas than VGO-L, differences in the RS reactor were negligible. Overall, for given feedstock, reaction temperature and conversion, the yields in the FFB reactor do not exceed those of the RS. As expected, the dry gas yields increased with reaction temperature for all the feedstocks and catalysts, but the sensitivity to this parameter was more significant in the RS reactor, with an about twofold increase against 30–50% increase in the FFB reactor for the 30 ◦ C delta. Both sets of data show that for the whole conversion ranges, catalyst E-cat L yields more dry gas than catalyst E-cat R, independently of the feedstocks and temperatures. The yields of LPG are shown in Fig. 5 as a function of conversion. It can be seen that the results in both reactors follow the same trend in the conversion scheme, with a stable primary plus secondary character for the group. For a given feedstock and catalyst, the yield curve from both reactors define the same line. The catalyst E-cat L yields more LPG than the catalyst E-cat R with the

same feedstock, which is tempting to be explained by the lower rare earth content of the former catalyst, inhibiting hydrogen transfer reactions and enhancing gasoline cracking; however, such difference in hydrogen transfer activity is not supported by the similar zeolite load and unit cell size measured for both catalysts. The paraffinic VGO-L yields more LPG than the aromatic VGO-R over the same catalyst. Moreover, the rate of LPG yield with conversion (slope of the yield curves) is systematically higher when cracking VGO-L, a dependency that may be related to the presence of ZSM-5 in both catalysts. It is well known in literature that the effectiveness of ZSM-5 in producing light olefins is the highest with the more paraffinic feedstocks [28,29]. The higher yields of LPG due to the 30 ◦ C increase in reaction temperature are revealed by the results from both reactors. Overall, and particularly in the case of the RS reactor, propylene, which is the product most favored by the addition of the ZSM-5 additive, showed yields with the same trends and dependences as LPG, so that the same discussion applies. In the case of

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

35

35

(b) 30

25

25

LPG, %wt

LPG, %wt

(a) 30

20

20

15

15

10 50

60

70

80

90

10 50

100

60

35

80

90

100

90

100

35

(c)

(d)

30

30

25

25

LPG. %wt

LPG, %wt

70

Conversion, %wt

Conversion, %wt

20

15

10 50

439

20

15

60

70

80

90

100

10 50

60

Conversion, %wt

70

80

Conversion, %wt

35

LPG, %wt

30

(e)

25

20

15

10 50

60

70

80

90

100

Conversion, %wt Fig. 5. Yield curves for LPG: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols as in Fig. 3.

the FFB reactor, the high conversions achieved may interfere with this view. Fig. 6 shows the yields of the gasoline cut (C5–C12). The results noticed in each reactor are different, since the yield curves observed in the RS reactor show that gasoline is a stable primary group, selectivities being about 54% at 500 ◦ C and 48% at 550 ◦ C, while the onset of an overcracking regime can be seen in the FFB reactor, even at the lower temperature of 510 ◦ C; moreover, in some cases, only the branch with negative slope as a function of conversion can be observed in the FFB reactor. In support of this observation, it should be noted that conversions were very high in the FFB. The setting up of an overcracking regime is possible in the range of reaction temperatures used in this work (see e.g., [30–32]). In the RS reactor, slightly more gasoline was obtained with the aromatic feedstock VGO-R than with the paraffinic VGO-L with both catalysts at every condition, while the opposite was true

in the FFB reactor, under overcracking regime. The comparison between catalysts’ performances shows very important differences in the FFB reactor: for given feedstock and conditions, the catalyst E-cat R clearly yields more gasoline than the catalyst E-cat L. These differences were quite smaller in the RS reactor, but led to the same catalyst ranking with both feedstocks and reaction temperatures. LCO yields are shown in Fig. 7. In the RS reactor the LCO from the aromatic feedstock VGO-R has a primary stable behavior, while that from the paraffinic feedstock VGO-L suggest to be under overcracking, particularly at 540 ◦ C; the yields from VGO-R are somewhat higher than those from VGO-L. Results in the FFB reactor show all the same behavior for LCO yields, which decrease as a function of conversion indicating the overcracking early at 510 ◦ C irrespective of the feedstock; in this reactor, clearly yields from VGO-R are higher than those from VGO-L. Both setups indicate that

440

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

50

50

(a)

(b) 45

Gasoline, %wt

Gasoline, %wt

45

40

35

30

40

35

30

25 50

60

70

80

90

25 50

100

60

Conversion, %wt 50

(c)

90

100

90

100

(d) 45

Gasoline, %wt

Gasoline, %wt

80

50

45

40

35

30

25 50

70

Conversion, %wt

40

35

30

60

70

80

90

100

25 50

60

Conversion, %wt

70

80

Conversion, %wt

50

Gasoline, %wt

45

(e)

40

35

30

25 50

60

70

80

90

100

Conversion, %wt Fig. 6. Yield curves for gasoline: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols as in Fig. 3.

catalyst E-cat R yields more LCO than catalyst E-cat L under all the conditions. 3.3. Product distributions It is important to know how products are distributed in the most important product groups in the experiments in each reactor. For example, olefinicity in the LPG group, calculated as the relationship between the yields of C3 and C4 olefins and the whole group (refer to Fig. 8), declines with conversion, more pronouncedly in the case of the FFB reactor, in coincidence with its prevailing overcracking regime. At constant conversion, LPG olefinicity slightly increased with the reaction temperature in the FFB reactor, corroborating the results of previous work [33]. Moreover, the paraffinic feedstock also led to higher LPG olefinicity

(see Fig. 8), which is in line with the previous comment about the higher potential for alkenes make by ZSM-5 additives with more paraffinic feedstocks [28]. On the other hand, despite its higher rare earth content, or even assuming the same hydrogen transfer ability in both catalysts (same UCS, same zeolite load), the catalyst E-cat R clearly showed increased LPG olefinicity as compared to the catalyst E-cat L. A possible explanation might be related to the remarkable differences in accessibility shown by the two catalysts (AAI of 15.0 for E-cat R and 4.5 for E-cat L): Lappas et al. [34] found that for a constant RE content, the higher accessibility catalysts assist for less hydrogen transfer reactions, thus inducing higher olefinicities. Associated to hydrogen transfer reactions, which consume olefins, the values of olefinicity observed in both reactors show that the catalyst E-cat L transfers more hydrogen than catalyst E-cat R, and that the higher temperature favors

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

20

20

(a)

18

16

16

14

14

LCO, %wt

LCO, %wt

18

12 10

10 8

6

6

4 50

60

70

80

90

100

(b)

12

8

4 50

60

Conversion, %wt

(c)

18 16

14

14

LCO, %wt

LCO, %wt

80

90

100

90

100

20

16

12 10

10 8

6

6 60

70

80

90

100

(d)

12

8

4 50

70

Conversion, %wt

20 18

441

4 50

60

Conversion, %wt

70

80

Conversion, %wt

20 18

(e)

LCO, %wt

16 14 12 10 8 6 4 50

60

70

80

90

100

Conversion, %wt Fig. 7. Yield curves for LCO: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols as in Fig. 3.

cracking reactions selectively due to their higher energies of activation. It is not possible to obtain fully consistent information about gasoline composition in the operation of the FFB unit because carry over to the gas product of most of the C5 compounds produced in the reactor is usually present. Moreover, some condensation in the gas sampling system may occur. While it is possible to include the C5 compounds in the gas phase back into gasoline composition, this approach may distort experimental information. The experiments in the RS reactor can provide a detailed analysis of the naphtha cut. In effect, the composition of the cut can be analyzed in terms of concentration of aromatics, naphthenics, n-paraffins, iparaffins and olefins, and the RON values can be assessed through Anderson-type [35] or proprietary correlations. However, only comparisons between the RS reactor and the DCR pilot unit can

be performed, since this information is not available in the FFB reactor. 3.4. Coke yields Coke yields deserve a special consideration due to the importance they have in the commercial process, where a delicate heat balance governs the overall operation. It had been shown that the RS reactor yields coke in amounts that are closer to the commercial operation than those from MAT reactors [22], as a direct consequence of the different contact modes in the reactors. Since the FFB reactor is similar to MAT-type reactors from the operative point of view, even though the actual contact between catalyst particles and reactants is completely different, it is expectable that results in the FFB reactor show the same tendency to high coke yields.

442

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

0.9

0.9

0.8

0.7

LPG Olefinicity

LPG olefinicity

0.8

(a)

0.6 0.5

0.7 0.6 0.5

0.4

0.4

0.3

0.3

0.2

50

60

70

80

90

100

(b)

0.2

50

60

0.8

LPG Olefinicity, wt%

LPG Olefinicity

(c)

0.7 0.6 0.5 0.4 0.3 0.2

50

80

90

100

90

100

0.9

0.9 0.8

70

Conversion, %wt

Conversion, %wt

(d)

0.7 0.6 0.5 0.4 0.3

60

70

80

90

100

0.2

50

60

Conversion, %wt 0.9 0.8

70

80

Conversion, %wt

(e)

LPG Olefinicity

0.7 0.6 0.5 0.4 0.3 0.2 50

60

70

80

90

100

Conversion, %wt Fig. 8. LPG olefinicity as a function of conversion: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols as in Fig. 3.

Coke yields are shown in Fig. 9, where it can be seen that there exist substantial differences between the laboratory reactors for a given set of catalyst, feedstock and reaction temperature. In the range of conversions obtained in the RS reactor, coke shows to be a stable primary product at all conditions (catalyst, feedstock and reaction temperature), while in the FFB reactor coke yields show steep positive slopes as a function of conversion. According to the previous discussion, yields in the FFB reactor are larger than those in the RS reactor. In the RS reactor, the coke yields obtained with the two feedstocks are very similar, or slightly higher in the case of the aromatic VGO-R. The FFB reactor showed a much higher sensitivity in relation to feedstocks, VGO-R yielding significantly more coke than the paraffinic VGO-L. The effect of the feedstock compositions on the coke yields is not so predictable in the present case, because both

the higher degree of aromaticity in VGO-R [36] (see Table 2) and the higher content of high boiling point hydrocarbons in VGO-L (as deduced from its distillation profile and RCR number) [22] can account for higher coke yields. Both reactors concur in catalyst E-cat L yielding more coke than catalyst E-cat R under similar conditions. In principle this is not expected, since FCC catalysts with lower matrix/zeolite ratio and rare earth content (less prone to promote hydrogen transfer reactions) tend to show better coke selectivity, producing less coke at constant conversion. However, it must be taken into account the considerable nickel contamination of E-cat L, which can induce a high dehydrogenation activity and consequently a high coke make. In all the cases the different reaction temperatures did not bring about significant changes in coke yields.

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

15

15

(a)

(b) 10

Coke, %wt

Coke, %wt

10

5

0 50

60

70

80

90

5

0 50

100

60

Conversion, %wt

80

90

100

90

100

15

(c)

(d)

10

10

Coke, %wt

Coke, %wt

70

Conversion, %wt

15

5

0 50

443

60

70

80

90

100

5

0 50

60

Conversion, %wt

70

80

Conversion, %wt

15

(e)

Coke, %wt

10

5

0 50

60

70

80

90

100

Conversion, %wt Fig. 9. Yield curves for coke: (a) RS 510 ◦ C, (b) FFB 510 ◦ C, (c) RS 540 ◦ C, (d) FFB 540 ◦ C, (e) DCR 540 ◦ C. Symbols as in Fig. 3.

3.5. Results in the DCR pilot plant One of the feedstocks (aromatic VGO-R) was tested over the two catalysts in a DCR circulating riser pilot plant at 540 ◦ C and two feed temperatures (two distinct C/O). Processing the other feedstock VGO-L led to reluctant operational and intermittent problems, so that only a single point with the catalyst E-cat R could be obtained for this feed with acceptable mass balance. The results are displayed in Figs. 3–9 along with the plots of the laboratory data. Conversion levels in pilot riser unit, ranging from 61 to 87%, were closer to the ones obtained in RS reactor, whereas FFB stayed over 80% at the same reaction temperature. It can be seen that conversions in the DCR unit were somewhat higher with catalyst E-cat R, similarly to the observations in the laboratory reactors (see Fig. 3).

However, the higher conversion noticed for the paraffinic feed in the pilot scale and with FFB testing could not be perceived in the RS reactor. In relation to dry gas yields, all the reactors showed similar values in the 2–5% range, and a positive slope as a function of conversion, which was very small in the case of the pilot unit. It is difficult, then, to define differences in the behaviors in each reactor. In the case of LPG, the yields observed in the laboratory reactors (very close between them) were higher than those in the DCR reactor for both catalysts and feedstocks; such difference between laboratory and DCR reactors is also found elsewhere [6], yet with no reasonable explanation for that. The higher dry gas and LPG yields for the catalyst E-cat L were evident in the laboratory reactors as well.

444

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

Table 3 Gasoline composition (wt%) at 75% conversion and calculated RON. RS and DCR reactors operated at 540 ◦ C with VGO-R and two types of catalyst. E-cat L

Aromatics n-Paraffins i-Paraffins Olefins Naphthenics RON

E-cat R

RS

DCR

RS

DCR

46.4 3.8 29.4 14.1 6.3 97.6

41.8 3.0 17.2 29.9 19.9 96.2

43.4 3.4 28.2 17.1 7.9 97.1

36.3 2.5 11.9 39.2 14.4 90.0

Gasoline yields showed a very near match between the RS reactor and the DCR pilot plant, for both catalysts, while the different yield-conversion trend in the FFB reactor is apparent. Gasoline composition in the results of the DCR unit (Fig. 10) can be compared only with that observed in the RS reactor. Table 3 shows the concentrations of the groups (naphthenics, olefins, nand iso-paraffins, and aromatics) in gasoline in each of the reactors. It can be seen that the proportions of aromatics and paraffins (iso and normal) in the gasoline produced in the RS laboratory reactor were always significantly higher than the corresponding concentrations in the DCR reactor, while the opposite can be observed for the concentrations of olefin and naphthenic groups, which were notably lower. These evidences, confirmed by the same type of conclusions derived from the comparison of olefinicity in LPG (Fig. 8), suggest that at a given conversion, hydrogen transfer reactions are more extensively developed in the laboratory reactor. Differences in effective contact time leading to time averaging effects [6] could be a reason for this. Table 3 also shows the calculated values of gasoline RON in both setups, which resulted similar in the case of E-cat L and diverged in the case of E-cat R; it must be considered that these values are calculated according to the composition of the gasoline cut. On the other hand, RS agreed with DCR as for the apparent evidence of higher hydrogen transfer activity in catalyst E-cat L (less olefins and naphthenics, more aromatics and paraffins). However, as aforementioned this is not compatible with the similar zeolitic characteristics of both catalysts. An alternative mechanism is likely prevailing in E-cat L and also favoring the consumption of olefins and naphthenics into aromatics; that could be the dehydrogenation effect promoted by the much higher concentration of nickel deposited on E-cat, which also explained its unexpected higher coke. LCO is the product group where most particular differences among reactors showed up (see Fig. 7). For example, for both catalysts, yields in the pilot plant were higher than those in the RS reactor (about 25–30% or more), and were about the same than those in the FFB reactor, but achieved at significant lower conversions. It was for the LCO, also, that the yield curves in each of the laboratory reactors showed a completely different behavior. Nevertheless, both laboratory reactors were in line with pilot riser unit with respect to catalyst ranking. In every case the catalyst E-cat R, with a higher matrix content than catalyst E-cat L, showed the expected higher LCO selectivity. The single point with VGO-L so close to the corresponding VGO-R curve may reflect no significant effect of the feedstock on the LCO selectivity, which is in line with the alike difficult discrimination in the lab reactors. Coke yields were similar in the pilot plant and RS reactors, matching both yield and conversion for both the two catalysts, probably due to the similar contact in these setups between an element of reactants and an element of catalyst particles, which ideally accompany each other during the contact time. Particularly for catalyst E-cat L, coke yields in the FFB reactor were very high.

Table 4 Results of the multiple regression analysis. Coefficients for each of the independent variables. Y

Int

Feed

Cat

RTx

CTO

R2

FFB reactor Dry gas LPG Gasoline LCO Bottoms Coke

−16.62 −17.07 56.06 34.21 42.37 4.46

−0.16 −7.26 −1.93 3.96 6.20 −0.82

−0.80 −2.63 7.04 3.49 −4.06 −3.04

0.04 0.08 −0.03 −0.04 −0.05 –

0.13 0.60 −0.31 −0.55 −0.95 1.08

0.97 0.98 0.95 0.94 0.88 0.97

RS reactor Dry gas LPG Gasoline LCO Bottoms Coke

−31.30 −28.72 29.52 7.52 113.58 −10.59

– −4.02 – – 3.78 –

−0.58 −3.04 3.56 2.99 – −0.81

0.06 0.09 – – −0.15 0.03

0.07 0.36 0.61 0.09 −1.21 0.10

0.91 0.90 0.72 0.49 0.77 0.77

Y = Int + A * Feed + B * Cat + C * RTx + D * CTO (or Time). E-cat R = 1, E-cat L = 0; VGO-R = 1, VGO-L = 0.

3.6. Statistical analysis and kinetic information In order to support the qualitative assessments, a statistical approach was used, based on multiple regression analysis. The regression analysis assumes a linear relationship between the experimental results (product yields) and independent variables (operating variables, feed and catalyst type), resulting in the equations displayed in Table 4. As independent variables the reaction temperature and the CTO (for the FFB reactor) and the reaction time (for the RS reactor) were used together with the feedstock and the catalyst types. Feed type and catalyst type were defined as dummy variables, assuming an arbitrary value of 1 or zero for each of the two possibilities. Bottoms yield can be interpreted as an inverse measure of conversion (100, bottoms) as defined previously. Then, the following observations could be drawn from the calculated coefficients (refer to Table 4, considering that empty entries mean that the effect is statistically not significant). (i) The RS reactor is not sensitive to feed quality, except for LPG and bottoms yields (see column Feed corresponding to RS in Table 4). (ii) The FFB reactor is operating in the overcracking regime, as indicated by the negative coefficient for C/O in the gasoline line. (iii) The response to reaction temperature in the RS reactor differs from that in the FFB reactor. Increasing the temperature in the FFB unit will decrease the yields of the liquid products (gasoline, LCO, bottoms) by about the same amount. Increasing temperature in the RS unit will have little effect on gasoline and LCO but a substantial effect on bottoms. (iv) The coefficients in column “Cat” are particularly important for catalyst ranking. There is good agreement between the RS and FFB units for dry gas, LPG, gasoline and LCO. The coefficients for coke and bottoms, however, are very different, and the reasons for this were discussed previously. If the information from the pilot unit is considered as a standard, then it can be concluded that both the RS and the FFB reactors will correctly predict the catalyst ranking for dry gas, LPG, gasoline and even LCO in spite of operating at different conversion levels. However, FFB unit will exaggerate the coke yield delta and the RS unit will not be able to easily realize the differences in bottoms yields (conversion). Even though it is not the main objective of this work, it is also possible to include some comparisons between the results obtained in the subsidiary modelling of the conversion of the feedstocks in

50

50

45

45

Aromatics in Gasoline, %wt

Olefins in Gasoline, %wt

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

40 35 30 25 20 15 10 5 0 50

60

70

80

90

40 35 30 25 20 15 10 5 0 50

100

60

70

25

25

20

20

15

10

5

60

70

80

80

90

100

90

100

Conversion, %wt

Paraffins in Gasoline, %wt

Naphthenics in Gasoline, %wt

Conversion, %wt

0 50

445

90

100

15

10

5

0 50

60

70

Conversion, %wt

80

Conversion, %wt

Fig. 10. Gasoline composition as a function of conversion in DCR tests at 540 ◦ C: (a) olefins, (b) aromatics, (c) naphthenics, (d) paraffins. Symbols as in Fig. 3.

K1

A1 K3

state assumption (the vapour phase sees catalyst of uniform age, the activity of the bed decaying uniformly). The RS was considered a batch, isothermal well stirred tank reactor. In both cases deactivation was considered to be first order. Mass balances were written accordingly and data fitting was performed with commercial software following the procedure by Gross et al. [43] in the case of the FFB reactor and a standard method in the case of the RS reactor. As expected, the kinetic constants derived from each laboratory reactor were very different. Results are shown in Table 5 with the examples of K0 (K1 + K3 , VGO conversion) and the deactivation parameter. It can be seen that for a given set of E-Cat-VGO-reaction temperature, the kinetic parameter is from 15 to 30 times larger in the case of the RS reactor. As expected, the ranking of K0 values reproduces the activity profiles shown in Fig. 3 and, in general, both the constants and the deactivation parameters increase with the reaction temperature. Some particularities can be observed, such as the inverse sequence of catalyst activities when the conversion

A2 K2

A3 Fig. 11. Three-lump kinetic model for VGO conversion [42].

both reactors. These two setups have been used to assess the kinetic constants from various models and reactions (see e.g., [37–41]). The very different fluidynamics conditions of the laboratory rectors should impact on the assessment of the kinetic constants; however, in order to do a simple comparison, the well known three-lump model (see Fig. 11, Weekman [42]) was used. The FFB reactor was considered isothermal, with plug flow of the gas phase, negligible interparticle diffusion and operating under the pseudo steady

Table 5 K0 = K1 + K3 kinetic constant in the three-lump model and ˛ deactivation parameter. T (◦ C)

Feed

Catalyst

K0 [h−1 ]

˛ [h−1 ]

FFB 510 540 510 540 510 540 510 540

VGO-L VGO-L VGO-L VGO-L VGO-R VGO-R VGO-R VGO-R

E-Cat L E-Cat L E-Cat R E-Cat R E-Cat L E-Cat L E-Cat R E-Cat R

35.335 71.111 49.809 62.976 38.677 45.253 33.180 35.335

RS ± ± ± ± ± ± ± ±

0.231 1.400 0.325 1.490 0.873 2.000 0.271 0.231

809.027 1149.784 854.244 1130.443 610.184 833.022 874.155 1024.406

K1 /K0

FFB ± ± ± ± ± ± ± ±

54.012 94.379 77.770 118.72 17.928 54.560 75.850 98.946

32.269 46.948 28.358 34.115 60.998 61.068 43.726 32.269

± ± ± ± ± ± ± ±

0.551 1.530 0.544 1.920 1.850 3.550 0.683 0.551

RS

FFB

RS

37.844 ± 64.218 53.324 ± 108.918 71.074 ± 78.480 171.444 ± 93.424 0 79.858 ± 54.322 203.213 ± 79.462 194.474 ± 78.418

0.62 0.55 0.70 0.77 0.65 0.64 0.83 0.85

0.56 0.50 0.62 0.58 0.59 0.54 0.65 0.63

446

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447

of VGO-R in both reactors is compared. Using the RS reactor de Lasa and Kraemer [44] converted a commercial VGO over an equilibrium catalyst in the range from 500 to 550 ◦ C, observing K0 values which are very close to those reported in Table 5. However, the first order deactivation parameters, which in these cases play a strong fitting role, were considerably larger (up to about three times) in their case. In the case of the FFB reactor, the constants reported by Gross et al. [43], who used commercial VGOs and catalyst, are in the same range as those in Table 5. A brief analysis of gasoline selectivity, defined by the K1 /K0 relationship, shows that it is larger in the FFB reactor than in the RS reactor, commercial units showing values closer to the second ones [45,46]. Using a FFB reactor, Gross et al. [43] also observed a high value of 0.81. For both reactors, the gasoline selectivity was higher in the case of converting VGO-R, a fact which is more noticeable for the gasoline yield curves in the case of the RS reactor (see Fig. 6). 4. Conclusions The catalytic performances of two different commercial equilibrium catalysts of the octane-barrel and resid types with two different commercial VGO feedstocks of the paraffinic and aromatic types were compared in two fluidized bed laboratory reactors, the FFB and the CREC Riser Simulator, and against results in a circulating pilot plant. Quantitative similarities between the laboratory reactors, which also included results in the pilot plant, were observed. Overall, the same catalyst ranking was observed in both laboratory reactors when the various product yields were considered, differences between catalysts and especially between feedstocks being more perceptible in the FFB reactor. Concerning the yields of the most important hydrocarbon groups, the CREC Riser Simulator reactor showed a linear yield curve for gasoline, which facilitated the analysis and comparisons; however, an overcracking regime was not shown, like in the case of the FFB reactor. The LCO yield curves were more defined and differences between catalysts appeared more clearly in the FFB reactor. Coke showed high yields in the FFB reactor, typical of confined beds with continuous feed of the reactants, similarly to what is observed in MAT-type reactors, while coke yields in the CREC Riser Simulator reactor were in the range of those observed in the pilot plant and commercial units. It can be concluded that the advantages shown by one of the laboratory reactors are the disadvantages shown by the other, thus showing a complementing potential in the laboratory evaluation of commercial catalysts and feedstocks. Acknowledgments Financial support from Universidad Nacional del Litoral, CAI+D 2009, Proy. 60-294; Agencia Nacional de Promoción Científica y Técnica PICT 2005 14-32930 and Consejo Nacional de Investigaciones Científicas y Técnicas (CONICET) PIP 1257/09 is gratefully acknowledged. The authors are also grateful to Petrobras S.A. for supporting the FFB experiments and the realization of DCR runs. References [1] P. O’Connor, Catalytic Cracking: the future of an evolving process, Stud. Surf. Sci. Catal. 166 (2007) 227–251. [2] R.H. Harding, A.W. Peters, J.R.D. Nace, New developments in FCC catalyst technology, Appl. Catal. A: Gen. 221 (2001) 389–396. [3] G.W. Young, Realistic assessment of FCC catalyst performance in the laboratory, Stud. Surf. Sci. Catal. 76 (1993) 257–292. [4] W.H. Humes, ARCO’s updated cat-cracking pilot unit, Chem. Eng. Prog 5 (February) (1983) 1–54. [5] U. Sedran, Laboratory testing of FCC catalysis and hydrogen transfer properties evaluation, Catal. Rev. Sci. Eng. 36 (1994) 405–431.

[6] D. Wallenstein, A. Haas, R.H. Harding, Latest developments in microactivity testing: influence of operational parameters on the performance of FCC catalysts, Appl. Catal. A: Gen. 203 (2000) 23–36. [7] L.T. Boock, X. Zhao, Recent advances in FCC catalyst evaluations: MAT vs. DCR pilot plant results, in: Symp. Advances in FCC Conversion Catalysts, Presented before the Division of Petroleum Chemistry, 211th National Meeting, ACSociety, New Orleans, LA, March 24–29, 1996. [8] E.L. Moorehead, J.B. McLean, W.A. Cronkright, Microactivity evaluation of FCC catalysts in the laboratory: principles, approaches and applications, Stud. Surf. Sci. Catal. 76 (1993) 223–255. [9] P. O’Connor, M.B. Hartkamp, A microscale simulation test for FCC development, Am. Chem. Soc. Div. Pet. Chem. 33 (4) (1989) 656–662. [10] R. Rawet, H.S. Cerqueira, J.C. Pinto, The influence of covariances during laboratory evaluation of FCC catalysts, Appl. Catal. A: Gen. 207 (2001) 199–209. [11] D. Wallenstein, M. Seese, X. Zhao, A novel selectivity test for the evaluation of FCC catalysts, Appl. Catal. A: Gen. 231 (2002) 227–242. [12] J. Biswas, I.E. Maxwell, Recent process- and catalyst-related developments in fluid catalytic cracking, Appl. Catal. 63 (1990) 197–258. [13] S.H. Ng, A. Humphries, C. Fairbridge, Y. Zhu, C. Khulbe, T.Y.R. Tsai, F. Ding, J.P. Charland, S. Yui, Comparisons of FCC product yields and qualities between reactors using Canadian heavy feeds, Fuel Proc. Technol. 86 (2005) 1335–1350. [14] H. I. de Lasa, Riser simulator, U.S. Patent 5,102,628 (1992). [15] G. de la Puente, U. Sedran, Evaluation of hydrogen transfer in FCC catalysts. A new approach for cyclohexene as a test reactant, Chem. Eng. Sci. 55 (2000) 759–765. [16] C. Bidabehere, U. Sedran, Simultaneous diffusion, adsorption, and reaction in fluid catalytic cracking catalysts, Ind. Eng. Chem. Res. 40 (2001) 530–535. [17] G. de la Puente, G. Chiovetta, U. Sedran, FCC operation with split feed injections, Ind. Eng. Chem. Res. 38 (1999) 368–372. [18] R. Spretz, U. Sedran, Operation of FCC with mixtures of regenerated and deactivated catalyst, Appl. Catal. A: Gen. 215 (2001) 199–209. [19] I. Tiscornia, G. de la Puente, U. Sedran, Recycling of low-value hydrocarbon cuts by means of multiple injections to FCC units, Ind. Eng. Chem. Res. 41 (24) (2002) 5976–5982. [20] F.J. Passamonti, G. de la Puente, U. Sedran, Laboratory evaluation of FCC commercial catalysts. Analysis of products of industrial importance, Catal. Today 133–135 (2008) 314–318. [21] D. Wallenstein, R.H. Harding, J. Witzler, X. Zhao, Rational assessment of FCC catalyst performance by utilization of micro-activity testing, Appl. Catal. A: Gen. 167 (1998) 141–155. [22] F.J. Passamonti, G. de la Puente, U. Sedran, Comparison between MAT flow fixed bed and batch fluidized bed reactors in the evaluation of FCC catalysts. 1. Conversion and yields of the main hydrocarbon groups, Energy Fuel 23 (2009) 1358–1363. [23] F.J. Passamonti, G. de la Puente, U. Sedran, Comparison between MAT flow fixed bed and batch fluidized bed reactors in the evaluation of FCC catalysts. 2. Naphtha composition, Energy Fuel 23 (2009) 3510–3516. [24] J. Kayser (Kayser Technology, US), Versatile fluidised bed reactor, PCT Int. Appl. WO98/52685, Application: PCT US98/11223, 26 November 1998. [25] F.J. Passamonti, G. de la Puente, U.A. Sedran, Reconversion of olefinic cuts from fluidized catalytic cracking naphthas, Ind. Eng. Chem. Res. 43 (2004) 1405–1410. [26] N. Ravichander, T. Chiranjeevi, D.T. Gokak, R.K. Voolapalli, N.V. Choudary, FCC catalyst and additive evaluation – a case study, Catal. Today 141 (2009) 115–119. [27] X. Dupain, M. Makkee, J. Moulijn, Optimal conditions in fluid catalytic cracking: a mechanistic approach, Appl. Catal. A: Gen. 297 (2005) 198–219. [28] Y.G. Adewuyi, D.J. Klocke, J.S. Buchanan, Effects of high-level additions of ZSM5 to a fluid catalytic cracking (FCC) RE-USY catalyst, Appl. Catal. A: Gen. 131 (1995) 121–133. [29] X. Zhao, R.H. Harding, ZSM-5 Additive in fluid catalytic cracking: 2. Effect of hydrogen transfer characteristics of the base cracking catalysts and feedstocks, Ind. Eng. Chem. Res. 38 (1999) 3854–3859. [30] T. Myrstad, H. Engan, Testing of resid FCC catalysts in MAT, Appl. Catal. A: Gen. 171 (1998) 161–165. [31] S.H. Ng, J. Wang, C. Fairbridge, Y. Zhu, L. Yang, F. Ding, S. Yui, Study of Canadian FCC feeds from various origins and treatments. 1. Ranking of fedstocks based on feed quality and product distribution, Energy Fuels 18 (2004) 160–171. [32] I. Torre, J.M. Arandes, M.J. Azkoiti, M. Olazar, J. Bilbao, Cracking of coker naphtha with gas-oil. Effect of HZSM-5 zeolite addition to the catalyst, Energy Fuels 21 (2007) 11–18. [33] W. Gilbert, C.A. Baptista, Andrea R. Pinho, Exploring FCC flexibility to produce mid-distillates and petrochemicals, Stud. Surf. Sci. Catal. 166 (2007) 31–39. [34] A.A. Lappas, Z.A. Tsagrasouli, I.A. Vasalos, A. Humphries, Effect of catalyst properties and feedstock composition on the evaluation of cracking catalysts, Stud. Surf. Sci. Catal. 134 (2001) 71–86. [35] P.C. Anderson, J.M. Sharkey, R.P. Walsh, Calculation of the research octane number of motor gasolines from gas chromatographic data and a new approach to motor gasoline quality control, J. Inst. Pet. 58 (560) (1972) 83. [36] R.H. Harding, X. Zhao, J. Deady, NPRA Annual Meeting, San Antonio, 1997. [37] A. Gianetto, H.I. Farag, A.P. Blasetti, H.I. De Lasa, Fluid catalytic cracking catalyst for reformulated gasolines. Kinetic modelling, Ind. Eng. Chem. Res. 33 (12) (1994) 3053–3062. [38] D. Kraemer, U. Sedran, H. de Lasa, Catalytic cracking kinetics in a novel riser simulator, Chem. Eng. Sci. 45 (1990) 2447–2452.

F. Passamonti et al. / Chemical Engineering Journal 183 (2012) 433–447 [39] M. Al-Sabawi, J.A. Atias, H. De Lasa, Kinetic modeling of catalytic cracking of gas oil feedstocks: reaction and diffusion phenomena, Ind. Eng. Chem. Res. 45 (5) (2006) 1583–1593. [40] T. Odedairo, S. Al-Khattaf, Kinetic analysis of benzene ethylation over ZSM-5 based catalyst in a fluidized-bed reactor, Chem. Eng. J. 157 (1) (2010) 204–215. [41] R.C. Vieira, J.C. Pinto, E.C. Biscaia Jr., C.M.L.A. Baptista, H.S. Cerqueira, Simulation of catalytic cracking in a fixed-fluidized-bed unit, Ind. Eng. Chem. Res. 43 (19) (2004) 6027–6034. [42] V.W. Weekman, Kinetics and dynamics of catalytic cracking selectivity in fixedbed reactors, Ind. Eng. Chem. Proc. Des. Dev. 8 (1969) 385–391. [43] B. Gross, D.M. Nace, S.E. Voltz, Application of a kinetic model for comparison of catalytic cracking in a fixed bed microreactor and a fluidized dense bed, Ind. Eng. Chem. Process Des. Develop. 13 (3) (1974) 199–203.

447

[44] H. de Lasa, D. Kraemer, in: de Lasa, H.I. y col. (Eds.), NATO-ASI Ser. E, vol. 225. Kluwer, The Netherlands, 1992, pp. 71–131. [45] S. Mandal, D. Bhattacharyya, V.B. Shende, A.K. Das, S. Ghosh, Impact of Additive Usage in Distillate Fluid Catalytic Cracking Operation, ACS Symposium Series 571, American Chemical Society, Washington, DC, 1994, pp. 335–348. [46] L.A. Casali, S.D. Rocha, M.L. Passos, R. Bastiani, R.D. Pimenta, H.S. Cerqueira, Equilibrium FCC catalyst performance simulation based on mixtures of hydrothermal deactivated samples, Stud. Surf. Sci. Catal. 166 (2007) 148–162. [47] A.K. Hakuli, P. Imhof, C.W. Kuehler, Proc. Akzo Nobel ECO-MAGIC Symp. Noordwijk, The Netherlands, June 10–13, 2001, p. F-4.