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CATALYTIC CRACKING OF ARAB SUPER LIGHT CRUDE OIL TO LIGHT OLEFINS: AN EXPERIMENTAL AND KINETIC STUDY Sulaiman Saleh Fahad Al-Khattaf, and Syed Ahmed Ali Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.7b04045 • Publication Date (Web): 19 Jan 2018 Downloaded from http://pubs.acs.org on January 20, 2018

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CATALYTIC CRACKING OF ARAB SUPER LIGHT CRUDE OIL TO LIGHT OLEFINS: AN EXPERIMENTAL AND KINETIC STUDY *

Sulaiman S. Al-Khattaf and Syed A. Ali Center of Research Excellence in Petroleum Refining & Petrochemicals, King Fahd University of Petroleum & Minerals, Dhahran 31261, Saudi Arabia ABSTRACT Catalytic cracking of Arab Super Light (ASL) crude oil (containing 46.1 wt.% naphtha-range fraction) was studied over zeolite Y- (Y-Cat) and MFI-based (Z-Cat) catalysts at 500-575 oC.

Experiments were conducted in a riser simulator by

varying the residence times from 1 to 10 s. ASL crude oil and the cracked products were divided into heavy fraction, naphtha and C1-C4 gases. Experimental results showed that additional naphtha is formed due to the cracking of heavy fraction and the formation of C1-C4 gaseous products occur mostly via cracking of naphtha. An increase in reaction time or temperature showed more pronounced effect on the propylene yield compared to that of ethylene. Z-Cat produced more ethylene and propylene, which was attributed to its higher acidity, shape selectivity and the higher hydrogen transfer reaction over Y-Cat. A three-lump model was appropriate for kinetic modeling of the catalytic cracking of ASL over Y-Cat. Comparison of the activation energies and rate constants showed that conversion of heavy fraction to naphtha (EHN = 9.89 kcal/mol) was easier compared to the cracking of naphtha to C1-C4 gaseous products (ENG = 15.79 kcal/mol). Direct cracking of heavy fraction to C1-C4 gaseous products was found to have highest activation energy (EHG = 79.89 kcal/mol) in the reaction scheme. Keywords: Catalytic Cracking, Crude Oil, Light Olefins, Kinetics *

Corresponding author: Dr. Sulaiman S. Al-Khattaf, Center of Research Excellence in Petroleum Refining and Petrochemicals, The Research Institute, King Fahd University of Petroleum and Minerals, Dhahran 31261, Saudi Arabia. Phone: +966 13 860 2029; email: [email protected]

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CATALYTIC CRACKING OF ARAB SUPER LIGHT CRUDE OIL TO LIGHT OLEFINS: AN EXPERIMENTAL AND KINETIC STUDY *

Sulaiman S. Al-Khattaf and Syed A. Ali Center of Research Excellence in Petroleum Refining & Petrochemicals, King Fahd University of Petroleum & Minerals, Dhahran 31261, Saudi Arabia

GRAPHICAL ABSTRACT

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CATALYTIC CRACKING OF ARAB SUPER LIGHT CRUDE OIL TO LIGHT OLEFINS: AN EXPERIMENTAL AND KINETIC STUDY Sulaiman S. Al-Khattaf*and Syed A. Ali Center of Research Excellence in Petroleum Refining & Petrochemicals, King Fahd University of Petroleum & Minerals, Dhahran 31261, Saudi Arabia

HIGHLIGHTS

Catalytic cracking of Arab Super Light (ASL) crude over Y- and MFI-based catalysts Conversion of heavy fraction to C1-C4 gases occurred via intermediate naphtha Three-lump model was appropriate for catalytic cracking of ASL over Y-Cat Heavy fraction conversion to naphtha was easier than naphtha to C1-C4 gases Direct cracking of heavy fraction to C1-C4 gases was the most difficult route

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1.

INTRODUCTION

Conventional crude oil refineries are generally geared towards production of transportation fuels while the petrochemicals produced from refineries are considered as by-products.

However, higher growth in demand for

petrochemicals such as light olefins (ethylene and propylene) compared to that of gasoline and diesel is compelling the refiners to look for alternate approaches to produce petrochemicals for enhancing profit margin [1,2]. Options available include integration of petrochemical production with refineries by including further downstream processing.

Direct conversion of crude oil – especially

super light or light crude oils – is gaining momentum due to its abundance and low cost.

While a lot of technological challenges exist in commercially

exploiting the direct conversion route for petrochemical production, the recent growth in publications and patents in this subject indicates that vigorous efforts are being made to overcome these hurdles [3-10]. The well-established route for the production of light olefins is the steam cracking of ethane, naphtha or gas oils [11]. However, the application of steam cracking process to whole crude oil is severely limited due to rapid fouling. Catalytic cracking provides a more flexible alternative to the steam cracking route with the possibility to widen the feedstock choice and alter the product slate. The fluid catalytic cracking (FCC) process, which is widely applied for conversion of low-value heavy feedstocks, offers an appropriate platform to investigate the direct conversion of crude oil. It should be noted that the FCC

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process was basically designed to process heavy feedstocks such as atmospheric residue and vacuum gas oil. The extent of light olefins production via FCC unit is reported to be highly dependent on the nature of feedstock and the type of catalyst [12,13].

Since crude oil contains a substantial amount of lighter

fractions (such as naphtha or gas oil), its catalytic cracking may not exhibit substantial conversion over conventional FCC catalysts [13,14]. Attempts were made to address this problem by modification of FCC catalysts, use of catalyst additives, operation at high-severity, or integration with other processes. Corma et al., recently reviewed the published research on the direct conversion of crude oil to olefins with a focus on application of FCC-type technologies [3]. Recently, we have reported the direct catalytic cracking of three types of crude oils to light olefins and naphtha in a microactivity test (MAT) unit over equilibrated FCC catalyst (E-Cat) mixed with MFI [15]. Results showed that the addition of MFI with varying Si/Al ratio to E-Cat increased the yield of light olefins. In another study, the product recycling was simulated by mixing light crude oil and low-value hydrocarbon streams [16]. The results indicated that such an approach can be a cost-effective means to increase the overall efficiency of the catalytic cracking process. This paper presents the results of an experimental and kinetic study conducted to investigate the catalytic cracking of Arab Super Light crude using catalysts containing USY and MFI zeolites. The study was conducted in a riser simulator reactor which allows very short contact time experiments. Product distribution,

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especially the yield of ethylene and propylene, provided an insight into the potential application of crude to chemicals. A three-lump model was satisfactorily applied to describe the kinetics of kinetic modeling.

2.

EXPERIMENTAL

2.1.

Materials. Arab Super Light (ASL) crude oil used in this study as a

feedstock for catalytic cracking was obtained from a local refinery. The properties of ASL are presented in Table 1. ASL is a very light (API gravity: 51o) and low-sulfur (1100 ppmw) crude oil having a naphtha (C5-220°C) content of 46.1 wt.%. The naphtha fraction is highly paraffinic (66 wt.%) while the naphthenic (19 wt.%) and aromatic (15 wt.%) contents are much less. ASL is a paraffinic crude oil having a K-factor of 12.55. These properties of ASL make it a suitable choice for direct conversion to chemicals via catalytic cracking. Two catalysts containing different zeolites (about 40 wt.%) were used for investigating the cracking of ASL: (i) Y-Cat which contains USY-zeolite; and (ii) Z-Cat which contains MFI zeolite. Y-Cat was an equilibrium FCC catalyst obtained from a refinery. Both catalysts contain alumina (20-30 wt.%) and kaolin (as a binder). A fresh sample of Z-Cat was steamed at 810oC for six hours to make it comparable to equilibrium FCC catalyst. The properties of these catalysts are presented in Table 2. The catalysts were calcined at 550°C (ramping rate 5°C/min) for 3 h prior to its performance evaluation.

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2.2. Catalysts Characterization. The textural properties were determined by nitrogen adsorption measurements at 77K, using Quantachrome Autosorb 1-C adsorption analyzer. Samples were outgassed at 220 °C under vacuum (10-5 Torr) for three hours before the physisorption of nitrogen. The specific surface area was determined from the adsorption data in the relative pressure (P/P0) range from 0.06-0.3, assuming the cross-section of the nitrogen molecule to be 0.164 nm2. Temperature-programmed desorption of ammonia (NH3-TPD) was carried out using Micromeritics Chemisorb 2750 equipped with a mass spectrometry detector. About 50 mg of the sample was preheated at 300 °C in a flow of helium (25 ml/min) for 2 h. This was followed by the adsorption of 10% NH3/He at 100 ºC for 30 min. Samples were then purged in a helium stream for 2 h at 120 ºC in order to remove physisorbed and H-bonded ammonia. Then, the samples were heated from 120 to 700 ºC at a heating rate of 10 ºC/min in a flow of helium (25 ml/min) while monitoring the desorbed ammonia using TCD. Powder X-ray diffraction (XRD) patterns were recorded on a Shimadzu powder diffraction system using CuKα radiation (λKα1 = 1.54051Å, 45 Kv and 35 mA). The XRD patterns were recorded in the static scanning mode from 5 - 60° (2θ) at a detector angular speed of 0.01 °/s and step size of 0.02°. The average crystallite sizes were calculated from the X-ray line broadening of the main diffraction peaks using Scherrer’s formula (Equation 1):

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where Bc (2θ) represents the peak width at half height after correction of the experimental broadening by the use of Warren `s equation (Equation 2):

where Be (2θ) and Br (2θ) represent the peak width at half height for the analyzed powder and silicon (standard material), respectively. 2.3

Catalytic Performance Evaluation. Catalytic cracking of ASL was

carried out in riser simulator reactor, a type of fluidized batch reactor, because of its suitability to conduct very short-contact time experiments. The design and operation of the reactor system was described previously [17,18]. Experiments were conducted at 500-575oC with a catalyst/reactant ratio of 5 (catalyst = 0.81 g, reactant feed = 0.162 g) for 1, 3, 5, 7 and 10 s each. A typical catalyst evaluation procedure was started with the test on fresh catalyst for 10 s. Then the catalyst was regenerated for 15 s by burning the coke before the next test. Such alternation continues till the tests for all the reaction times at a particular temperature were completed. A fresh sample of catalyst was used for testing at next temperature. In this manner, the simple conversion was determined and catalyst deactivation during the test was minimized. Some experiments were repeated to check for reproducibility, which was found to be within a range of ± 2%. Reaction products were analyzed using an Agilent 6890N GC equipped with

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a flame ionization detector (FID) by means of a 60 m capillary column INNOWAX, cross-linked polyethylene glycol with an internal diameter of 0.32 mm. Carbon deposited on the spent catalyst was determined using a Horiba analyzer (EMIA). About one gram of the spent catalyst along with tungsten (combustion promoter) was burnt in the high temperature furnace. The CO2 thus formed was passed through an infra-red analyzer to calculate the carbon content in the sample. 3.

RESULTS AND DISCUSSION

3.1.

Physicochemical Characteristics. The properties of the catalysts are

presented in Table 2. As expected, the Y-Cat, which was based on Y-zeolite, possess pores with higher average diameter (10.9 nm) compared to Z-Cat (7.4 nm) which contains MFI zeolite. Pore volume of Y-Cat was found to be twice that of Z-Cat while the surface area of Y-Cat was higher than that of Z-Cat. These differences in the textural properties of the catalysts are expected to play a role in the catalytic cracking of ASL. NH3-TPD was used to determine the acidity of the catalysts. The total acidity of Z-Cat (Table 2) was found to be more than five times higher than that of Y-Cat. However, most of the acidity of Z-Cat was weak or moderate acidity which was recorded at temperatures of less than 450 oC. There was no difference in the strong acidity recorded at temperatures higher than 450 oC. These disparities in

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the acidity of different strengths between the catalysts are expected to influence the cracking activities. Powder XRD patterns of Y-Cat and Z-Cat are presented in Figure 1. In case of Y-Cat, an intense peak at 6.0° (2Ɵ) showed the presence of (1,1,1) crystal plane along with other peaks at around 2θ = 10.0°, 11.9°, 15.6°, 18.7°, 20.4°, 23.8° and 27.1° which are attributed to (220), (311), (331), (333), (440), (533), and (642) crystal planes, respectively. This peaks indicate the typical XRD patterns Y zeolite as reported in literature [19]. For Z-Cat, an intense peak at 22-25° along with peaks between 7-9° showed the characteristic peaks of MFI [20]. 3.2

Catalytic Cracking of ASL. Catalytic cracking of ASL was carried out

over Y-Cat and Z-Cat at 500-575 oC. Experiments were carried out by varying the residence times from 1 to 10 s and the product samples were analyzed using GC. For evaluation purpose, the products composition was divided into three components: (i) heavy fraction; (ii) naphtha fraction; and (iii) C1-C4 gases. Experimental values of these components obtained at 500, 525, 550, and 575 oC over Y-Cat and Z-Cat are presented in Figures 2 and 3, respectively. Product distribution obtained over Y-Cat showed a steady increase in conversion of heavy fraction and formation of C1-C4 gases with temperature as well as residence time. However, the naphtha fraction increased initially (at 1 s) and then decreases as the residence time increases. Similar trends were obtained over Z-Cat, with the exception of absence of initial increase in naphtha fraction

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at 575oC. At lower temperatures, the differences in product distribution obtained over the two catalysts were within a narrow range of 2-3 wt.% as summarized in Table 3. In fact, the difference in the behavior of two catalysts was evident only at 575oC with Z-Cat exhibiting somewhat higher conversion of heavy fraction and naphtha. At the shortest residence time (1 s), it was observed that the heavy fraction was converted mostly to naphtha. Hence the amount of naphtha was more than in the feed (46.1 wt.%). As the reaction time increases, the formation of C1-C4 gaseous products picked up along with the simultaneous decrease in the heavy fraction and naphtha contents. These trends suggest that the additional naphtha is formed from the cracking of heavy fractions and the formation of C1-C4 gaseous products occur mostly via cracking of naphtha. Compared to the decrease in heavy fraction and increase in C1-C4 gases up to the residence time of 10 s, the change in the content of naphtha fraction was modest. For example, product distribution over Y-Cat at 525 oC and 10 s indicate that the heavy fraction decreased by 22.3 wt.% and the C1-C4 gases increased by 18.6 wt.% while the increase in naphtha fraction was only 3.6 wt.% (Table 3). These trends also indicate that naphtha fraction was acting like an intermediate during the catalytic cracking of heavy fraction of ASL to form C1-C4 gases. The coke laydown after a 10s experimental run at 550oC was measured. The amounts of carbon found was 1.71 and 0.63 mgC/g.cat on Y-Cat and Z-Cat,

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respectively. The three-fold higher coke laydown on Y-Cat clearly indicates that cracking and/or dehydrogenation reactions involving cyclic molecules took place inside the pores of Y-Cat leading to the formation of coke precursors [21]. Such reactions are not expected to take place heavily over Z-Cat due to limited volume inside its narrow pores. The shape selective characteristic of Z-Cat enhances the reactions involving straight chain molecules such as cracking of higher alkanes via β-scission resulting in higher yield of light olefins. These differences in catalytic structure and reactions cause significantly higher coke laydown and lower light olefins yield over Y-Cat as compared to Z-Cat [22]. 3.3

Formation of Light Olefins. As discussed in the introduction, formation

of light olefins – especially ethylene and propylene – is the one the main objectives of catalytic cracking of ASL. The yields of ethylene and propylene obtained at 500, 525, 550, and 575 oC over Y-Cat and Z-Cat are presented in Figures 4 and 5, respectively. The experimental data showed that increase in reaction time has more pronounced effect on the increase in propylene yield compared to that of ethylene. Similar trend was observed with the increase in temperature as shown in Figure 6. With the increase in reaction temperature from 500 to 575 oC, the increase in ethylene yield over Y-Cat was only 3.4 wt.% whereas the propylene yield increased by 8.1 wt.%. Generally, Z-Cat produced higher yields of ethylene compared to Y-Cat at the same temperature. Since it was difficult to define and calculate the conversion due to simultaneous formation and cracking of naphtha fraction, the yield of C1-C4 gases was

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considered as a measure of the extent of reaction. A comparison of the light olefins yield as a function of C1-C4 yield at 550oC and 10s (Figure 7) also showed that Z-Cat produces more olefins than Y-Cat.

The yields of ethylene

and propylene were about 1 wt.% and 2. wt.% higher over Z-Cat at these conditions. These results can be attributed to the higher acidity and shape selectivity of Z-Cat. The propylene/ethylene (P/E) ratio indicates the selectivity to produce propylene which is dependent on the feedstock type and cracking process conditions. Typical P/E ratio of olefins produced by steam cracking of vacuum gas oil (VGO) was reported to be about 0.65 whereas the conventional catalytic cracking of VGO increases the P/E ratio to 2-3. This shows that ethylene formation is generally favored during thermal cracking while propylene formation is enhanced by catalytic cracking [2]. The results from catalytic cracking of ASL showed that the P/E ratio generally increased with residence time but decreased with temperature for the Y-Cat (Figure 4). At lower temperatures (500-550 oC), the value of P/E ratio was 4-6 and it decreased to about 2 at 575oC. It should be noted that propylene was catalytically produced by β-scission of longer chain alkanes [22]. Since catalytic cracking is predominant at lower temperatures (500-550oC), higher amount of propylene is produced compared to ethylene. As the temperature is increased to 575 oC, more ethylene is produced as the share of thermal cracking increases resulting in lower P/E ratio. The P/E ratio obtained over Z-Cat were in the range

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of 2-3 and generally decreased with residence time. As mentioned earlier, Z-Cat produced more olefins due to its shape selective property. Moreover, higher residence time allowed formation of higher amounts of ethylene (possibly due to secondary cracking) thus causing a decrease in P/E ratio. This difference in P/E ratio between the two catalysts indicate that Z-Cat favored the formation of more ethylene compared to Y-Cat, which could be attributed to the higher acidity and lower pore size of Z-Cat. The narrow pores of MFI enhances shape selective cracking of longer chain paraffins to lower olefins [23]. The occurrence of hydrogen transfer during FCC process can alter the yield of light olefins. Hydrogen transfer is a bimolecular reaction in which naphthenes donate hydrogen to olefins to form aromatics while the olefins are hydrogenated to paraffins. The extent of hydrogen transfer reaction can be estimated by defining a hydrogen transfer index as the sum of C2-C4 paraffins divided by the sum of C2-C4 olefins [24]. The effect of the extent of reaction on the hydrogen transfer index is presented in Figures 8(a) and 8(b) for Y-Cat and Z-Cat, respectively.

It can be noticed that the index generally decreases with

temperature due to the formation of more olefins accompanied by their lower adsorption on the catalyst surface. However, as the reaction time increases, there is a tendency towards an increase in the index indicating the occurrence of hydrogen transfer reactions. The effect is more obvious in Y-Cat than Z-Cat – especially at higher temperature – because the bimodal hydrogen transfer reaction requires larger pores to take place which is facilitated by Y-zeolite and

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not by MFI. As shown in Figure 9, the hydrogen transfer index at the same gas yield over Y-Cat is much higher than over Z-Cat at both conditions. With the increase in temperature and gas yield, the decrease in index over Z-Cat was significant due to production of more olefins and absence of hydrogen transfer reaction. 4.0

KINETIC MODELING

Crude oils contain thousands of chemical species with a wide distribution of boiling temperatures. Therefore, kinetic modeling of crude oil catalytic cracking can be carried out using a lumping strategy in which chemical species with similar properties are grouped together forming a smaller number of “pseudo” species [25]. A three-lump model consists of heavy fraction, naphtha and C1-C4 gaseous products can be considered as one of the simplest kinetic models for crude oil cracking. The heavy fraction lump (H) of crude oil is defined as the fraction boiling above 220oC; naphtha lump (N) contains the fraction between C5 and the hydrocarbons with a boiling point up to 220oC; and gaseous lump (G) consists of C1-C4 hydrocarbon gaseous products. The formation of coke is neglected from the reaction network because the rate of formation of coke is very low compared to other reactions [26]. However, a catalyst deactivation function was used to account for decrease in catalyst activity in the kinetic equations. A graphical representation of the proposed three-lump model for the kinetic model of catalytic cracking of ASL is presented in Figure 10.

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Modeling of the product distribution data obtained over Y-Cat was carried out to investigate of the kinetics of the main reactions that occur during the catalytic cracking of ASL. A simplified reaction network involving the following three key reactions is proposed and used for kinetic modeling: Reaction 1: Conversion of Heavy Fraction to produce Naphtha; Reaction 2: Conversion of Naphtha to produce C1-C4 Gaseous Products; and Reaction 3: Conversion of Naphtha to produce C1-C4 Gaseous Products. Additional assumptions for the kinetic model were: a)

The reactions are elementary;

b)

A single catalyst deactivation function defined for all the reactions taking place; and

c)

The reactor operates under isothermal conditions, which is justified by the negligible temperature change observed during the reactions.

The riser simulator, which is a fluidized batch bed reactor, was used for the reaction. For each independent reacting species, the material balance equation was: ܸ ݀‫݅ܥ‬ = ‫߮ ݅ݎ‬ ܹ‫ݐ݀ ܥ‬

(3)

where ‫ݎ‬௜ and ‫ܥ‬௜ represent the reaction rate and mole concentration of each species in the system, V is the volume of the reactor, ܹ஼ is the weight of catalyst, t is time in seconds, while ߮ is the deactivation function which accounts for

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decrease in the activity of the catalyst. The reaction order assigned to the three lumps was based on the reactivity of different pseudo-species. Since the heavy fraction contains a mixture of a very large number of compounds of widely different properties, it was suggested that a second order should be assigned to the cracking of heavier fractions. Due to a rather limited range of hydrocarbons (C5 to C12), a first order was generally used for naphtha cracking [27,28]. Based on the simplified reaction scheme (3-lump model) and considering the above mentioned assumptions and practices, the following set of differential equations are formulated to describe the molar balance of the species in the reaction medium during the catalytic cracking of ASL: (i)

Rate of disappearance of heavy fraction: V dC H = − k HN C H2 − k HG C H2 φ WC dt

(

(ii)

)

Rate of formation and disappearance of naphtha: V dC N = k HN C H2 − k NG C N φ WC dt

(

(iii)

(4)

)

(5)

Rate of formation of C1-C4 gaseous products: V dCC = k HG C H2 + k NG C N φ WC dt

(

)

(6)

where CH, CN, and CG are the molar concentrations of heavy fraction, naphtha, C1-C4 gaseous products in the reactor, respectively. V is the volume of the riser

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simulator reactor (45 cm3), WC is the mass of the catalyst (0.81 g) and t is the reaction time in seconds.

The reaction rate constants kHN, kNG, and kHG

correspond to the reactions 1, 2, and 3, respectively. The quantity φ is the catalyst deactivation function, which is defined as φ = e−λ x

(7)

H

where λ is catalyst decay constant based on conversion of heavy fraction. ki is the temperature dependent rate constant given by the Arrhenius relation below: ݇݅ = ‫݁ ݅ܣ‬

−‫݅ܧ‬ൗ ܴܶ

(8)

In order to reduce parameter interaction between the pre-exponential factor (Ai) and activation energy (Ei), a re-parameterization of ki by entering the temperature at an average reaction temperature of To, i.e. −‫ ݅ܧ‬1 1 ݇݅ = ݇‫ ݌ݔ݁ ݋‬൤ ൬ − ൰൨ ܴ ܶ ܶ0

(9)

where ko is the rate constant at To. Also, the concentration, Ci was expressed in terms of mass fraction of each species yi, which are the measurable variables from the chromatographic analysis, i.e. ‫= ݅ܥ‬

‫ܹ ݅ݕ‬ℎܿ ܸ‫ܹ݅ܯ‬

(10)

where Whc weight of liquid feedstock injected into the reactor (0.162 g), MWi = molecular weight of species i in the system (MWA = 330 g/mol, MWB = 115 g/mol; MWC = 48 g/mol); and V was the volume of the riser simulator reactor

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(45 cm3). 0 0 0 The kinetic parameters k HN , k NG , k HG , E HN , E NG , E HG and α were determined by

nonlinear regression analysis using MATLAB® package lsqcurvefit. Values of kinetic parameters and their associated 95% confidence limits as obtained from the regression analysis are presented in Table 4. The rate constant for conversion of heavy fraction to naphtha (kHN) was significantly higher than the rate constant for conversion of heavy fraction to C1C4 gases (kHG) – especially at 500-550oC. However, with the increase in temperature, kHG increased much faster than kHN and hence they become comparable at 550oC. Higher conversion rate of heavy fraction to naphtha compared to the gaseous products was in agreement with the values reported by other researchers for the kinetic modeling results of catalytic cracking of vacuum gas oil (VGO) [26,29,30]. Ancheyta-Juarez et al. [31] reported that kinetic rate constants for the conversion of VGO, which can be considered as a heavy fraction, to gasoline at 500oC was six times higher compared to the rate constant for the conversion of VGO to C3-C4 gases. The estimated activation energies for the conversion of heavy fraction to naphtha (EHN) and conversion of naphtha to C1-C4 gaseous products (ENG) were 9.89 kcal/mol and 15.79 kcal/mol, respectively.

Reported values for the

corresponding reactions during catalytic cracking of VGO were in the range of 11.7-16.2 kcal/mol and 13.6-27.0 kcal/mol, respectively [26,29-32]. It can be

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noticed that the ENG obtained in this study was within the range of activation energies reported in the literature while EHN is somewhat below the range. Comparison of the activation energies and rate constants showed that conversion of heavy fraction to naphtha was relatively easy compared to the cracking of naphtha to C1-C4 gaseous products. Direct cracking of heavy fraction to C1-C4 gaseous products required much higher activation energy (79.89 kcal/mol). Such a high value for EHG can be attributed to the presence of large quantity of naphtha in the reaction atmosphere which hinder the heavy fraction from adsorption and taking part in the reaction – especially at lower temperatures (500-550oC). Thus competitive reactions resulted in much lower conversion of heavy fraction directly to C1-C4 gaseous products. These results showed that the formation of C1-C4 gaseous products from heavy fraction mostly occurred via formation of naphtha (as an intermediate), followed by secondary cracking of naphtha to gaseous products. The Arrhenius plot, presented in Figure 11 showed that the rate of direct conversion of heavy fraction to gaseous products is highly dependent on the reaction temperature. At high temperature (575oC), the rate of direct conversion was substantial and compared with the rate of conversion of heavy fraction to naphtha. Graphical comparisons between the experimental data and model predictions based on the optimized parameters for the kinetic model are shown in Figure 12. It can be seen that there is excellent match between the experimental values and

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the model predictions for C1-C4 gaseous products whereas the correspondence in the predicted and experimental values of heavy fraction and naphtha is good. 5. CONCLUSIONS This study showed that direct catalytic cracking of ASL crude oil can be an alternate route to produce petrochemicals – especially propylene and ethylene – which are in great demand. Experimental results showed that that the formation of naphtha occurred from the cracking of heavy fractions and the formation of C1-C4 gaseous products occur mostly via cracking of naphtha. Hence, only a modest decrease in the yield of naphtha was observed under the conditions studies. An increase in reaction time or temperature showed more pronounced effect on the propylene yield compared to that of ethylene. Z-Cat produced more ethylene and propylene at the same yield of gaseous products as Y-Cat at all the conditions studied. These results are supported by higher hydrogen transfer index and higher coke laydown over Y-Cat. Shape selectivity of MFI also played a role in producing higher amounts of ethylene over Z-Cat. A three-lump model appropriately described the kinetics of catalytic cracking of ASL over Y-Cat. Comparison of the activation energies and rate constants showed that conversion of heavy fraction to naphtha (EHN = 9.89 kcal/mol) was easier compared to the cracking of naphtha to C1-C4 gaseous products (EHN = 15.79 kcal/mol) whereas the direct cracking of heavy fraction to C1-C4 gaseous products required much higher activation energy (EHG = 79.89 kcal/mol). It seems that the formation of C1-C4 gases mostly occurred via formation of naphtha as an intermediate. However, the rate of direct conversion of heavy fraction to gaseous products is highly dependent on reaction temperature and it

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may become significant above 575oC.

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Energy & Fuels

ACKNOWLEDGEMENTS The authors wish to acknowledge the support of the Research Institute of King Fahd University of Petroleum and Minerals (KFUPM). We also acknowledge the Ministry of Education, Saudi Arabia for establishing the Center of Research Excellence in Petroleum Refining and Petrochemicals at KFUPM. The support of Dr. Michael Klein, Saudi Aramco Chair Professor at KFUPM, and Dr. Craig Bennet in kinetic modeling is highly appreciated. The contribution of Late Mr. Mariano Gica through conducting the Riser Simulator experiments is gratefully recognized and acknowledged. NOMENCLATURE Symbols

Cj

concentration of species or component j in the reactor (mol/m3)

Ea

apparent activation energy of cracking of species a (kJ/mol)

Ka

kinetic rate constant of cracking of species a (m3/(kg cat s))

t

reaction time (s)

V

volume of the reactor

Wc

mass of the catalyst

X

Conversion of reactant

Greek letters

α

catalyst decay constant based on reactant conversion

φ

catalyst deactivation function

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REFERENCES [1]

Special Report: Asia petrochemical outlook H1 2017 olefins & polymers, Jan 2017, www.platts.com/petrochemicals.

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Powers, D., US Patent 7019187, Mar 28, 2006 (Equistar)

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Corma, A.; Sauvanaud, L.; Mathieu, Y.; Al-Bogami, S.; Bourane, A.; AlGhrami, M. Fuel, 2018, 211, 726-736.

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Khare, G.; Arne, M. Steam cracking of crude oil, IHS Chemical PEP Report 29J, Dec. 2015.

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Energy & Fuels

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Usman, A.; Siddiqui, M.A.B.; Hussain, A.; Aitani, A.; Al-Khattaf, S. Chem. Eng. Res. Des., 2017, 120, 121-137.

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Treacy, M.M.J.; Higgins, J.B.; Ballmoos, R. Collection of simulated XRD powder diffraction patterns for zeolites, 3rd revised edition, London: Elsevier; 1996.

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Energy & Fuels

Table 1.

Properties of Arab Super Light Crude Oil.

Property

Value

Gravity, °API

51.3

Density at 15°C, kg/m3

774

Sulfur (wt.%)

0.11

Vanadium (ppm)

1

Nickel (ppm)

Catalytic Cracking of Arab Super Light Crude Oil to ... - ACS Publications

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