Hydrocracking of Athabasca VR Using NiO-WO3 Zeolite-Based

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Hydrocracking of Athabasca VR using NiO-WO3 Zeolite Based Catalysts Thomas Kaminski, Shaheen Anis, Maen M. Husein, and Raed Hashaikeh Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.7b03754 • Publication Date (Web): 24 Jan 2018 Downloaded from http://pubs.acs.org on January 25, 2018

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Hydrocracking of Athabasca VR using NiO-WO3 Zeolite Based Catalysts Thomas Kaminski1, Shaheen Fatima Anis2, Maen M. Husein1,*, Raed Hashaikeh2,* 1

Department of Chemical & Petroleum Engineering, University of Calgary, Calgary, AB, Canada T2N 1N4 2 Chemical Engineering Department, Khalifa University of Science and Technology, Masdar Institute, Masdar City, P.O. Box 54224, Abu Dhabi, United Arab Emirates. ABSTRACT Hydrocracking of Athabasca vacuum residue (AVR) was carried out in an autoclave using particle and fiber forms of NiO-WO3 zeolite-supported catalyst. AVR hydrocracking was performed at 400°C at low and high H2 pressure of 70 psi and 365 psi, together with the corresponding control thermal cracking runs. The yield of the different products and the quality of the upgraded liquid was used to assess the catalyst performance. Similarity among energy consumption for the different samples suggested major thermal cracking endothermic reactions. In general, the catalytic runs provided better quality maltene product, whereas better quality product oil was only attained at high pressure. The catalytic runs at low H2 pressure gave the highest yield of combined asphaltenes and toluene insolubles. This yield, on the other hand, was the lowest for the fiber form at high H2 pressure. Simulated distillation results captured the superior performance of the fiber catalyst at high H2 pressure and showed ~50% conversion of the residue. On the other hand, the zeolite particles showed poor performance at high pressure with only ~30% residue conversion.

Keywords: thermal cracking; hydrocracking; asphaltene; maltene; heavy oil; zeolite *Corresponding author: Department of Chemical & Petroleum Engineering, University of Calgary, Calgary, AB, Canada T2N 1N4, Tel: (403) 220-6691; Fax: (403) 282-3945; E-mail: [email protected]; OR Chemical Engineering Department, Khalifa University of Science and Technology, Masdar Institute, Masdar City, Abu Dhabi, 54224, United Arab Emirates. E-mail address: [email protected].

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List of abbreviations Abbreviation

Definition

AVR

Athabasca vacuum residue

BAS

Bronsted acid site

HDC

Hydrocracking

LAS

Lewis acid side

TC

Thermal cracking

TI

Toluene insolubles

1. INTRODUCTION The demand for liquid fuels has increased in the recent years, creating a pressing need to explore sources other than conventional oil resources1. Heavy, unconventional oils constitute the majority of the current oil reserves, however due to the presence of heavy molecules such as asphaltenes, they are difficult to handle and transport and must be upgraded into lighter fractions2. Many upgrading processes are currently in use by industry, such as vis-breaking, TC and HDC, with the general objective of maximizing the liquid yield and enhancing its quality3. HDC is an upgrading process which breaks down long hydrocarbon chains into smaller, more useful hydrocarbons with higher a H:C ratio4. Compared to other upgrading processes, HDC is known to produce a higher quality and more stable product with less coke production from a wider range of feedstock5. While HDC is a relatively high temperature and pressure process, its operating conditions can be made less severe in presence of a catalyst6. HDC catalyst is a bifunctional catalyst which comprises of hydrogenation/dehydrogenation elements such as Ni-W 7, Ni-Mo8 or Co-Mo9 and an acidic support such as zeolite to crack the feed molecules10. The distance between the metallic and acidic sites on a catalyst largely impacts HDC performance,


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since the HDC reactions proceed in steps from site to site11. A typical reaction scheme for HDC can be described as follows. Initial dehydrogenation of the molecule on a metallic site, transport of the dehydrogenated molecule to an acidic site, cracking on the acidic site, transport to a metallic site, and finally hydrogenation of the cracked shorter chain molecules12. Numerous studies have reported high HDC activity of Ni-W hydrogenation elements on a zeolite support7– 13

. There has been a growing interest in fabricating catalyst in the fiber form as opposed to

particles, since fibers are less susceptible to agglomeration/aggregation. As such, fibers furnish more accessible active sites compared with particles13. Recently, HDC fibers made from NiOWO3-Zeolite Y were fabricated using the electrospinning technique14. In brief, this method utilizes an electrospinning suspension containing the desired elements such as the metal salts and zeolite dispersed in a solvent such as water and/or ethanol. Usually a polymer is added to provide the right viscosity necessary for electrospinning15. The electrospinning suspension is filled in a syringe whose needle is connected to a collector through high voltage supply. The nozzle of the needle is fixed at a certain distance from the collector. On application of a high voltage, the liquid droplet at the tip of the needle becomes charged, and emerges as a fine jet which gets collected on the collector16. A uniform metal distribution within the fibers was achieved through this method. This benefits HDC reactions as it provides shorter diffusion path for the reacting molecules between the acidic and the hydrogenation sites17. The reported BET surface area of the HDC fibers was 309 m2/g, higher than that of the particles, 275 m2/g16. This study focuses on the HDC of Athabasca vacuum residue (AVR) and investigates the effect of the morphology of NiO-WO3 on zeolite catalyst on HDC reactions. While HDC has been the topic of interest for some time, the effects of the catalyst and its properties have yet to be fully investigated. Zeolite catalytic fibers have shown potential in various cracking reactions13. Nevertheless, zeolite fibers for hydrocracking is still a new research area where only recently zeolite Y-NiO-WO3 fibers were fabricated and tested for simple n-heptane feed molecule


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reaction14. Thus, exploring the potential use of zeolite Y-NiO-WO3 fibers as a catalyst for heavy feedstock upgrading opens up new research perspectives, which to the best of our knowledge has not been investigated yet. In this work, two zeolite catalyst morphologies are compared, fiber and particle forms. The liquid yield and its quality in terms of viscosity and °API gravity were used as criteria to evaluate catalyst performance18. Control runs involving non-catalytic HDC as well as TC were used for comparison in order to better isolate the role of the catalyst18. Energy consumption was recorded during the upgrading reactions in order to further characterize the types of reactions occurring at various conditions. Such an analysis, while rarely done, can be very beneficial to understanding heavy oil upgrading18. 2. EXPERIMENTAL PROCEDURE 2.1 Materials Zeolite-Y (CBV 720, Si/Al ratio of 30) was purchased from Zeolyst International (Conshohocken, PA, USA). Polyvinylpyrrolidone (PVP, MW = 1,300,000), ammonium metatungstate hydrate (AMT, >85 wt% WO3), Nickel (II) acetate tetrahydrate (NiAc, 98% pure), and ethanol were purchased from Sigma-Aldrich (MO, USA). Hydrogen (99% pure, Praxair Specialty Gas & Equipment, AB, Canada) was used to purge and pressurize the reactor, while nitrogen (99% pure, Praxair Specialty Gas & Equipment, AB, Canada) was used to purge the reactor for the TC control runs. Following upgrading, toluene insolubles (TI) and asphaltenes were rejected from the oil using toluene (BDH 99.8%, VWR, Canada) and n-heptane (BDH technical, VWR, Canada), respectively, as detailed below. 2.2 Catalyst Preparation and Characterization In this study, two Ni-W zeolite-based catalysts having particle and fiber morphologies were used. The catalysts were prepared according to the method reported by Anis et al.19. In brief, nickel and tungsten salts were impregnated 20 onto zeolite to obtain a Ni:W atomic ratio of 1:321. 4 ACS Paragon Plus Environment

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The fiber catalyst was prepared through electrospinning method, where the NiAc-AMT-Zeolite suspension in deionized (DI) water was electrospun in the presence of PVP employing the following process parameters14. solution flow rate of 0.5 mL/h, electrospinning voltage of 28 kV and an electropsinning distance of 15 cm. The post treatment of both the particles and the fiber catalysts consisted of overnight drying in a furnace at 80°C to remove the moisture, and catalyst calcination at 550°C at a heating rate of 1°C/min for 2 h under ambient conditions. Calcination decomposes the salt to yield the zeolite-NiO-WO3 catalyst, while it also removes the polymer. The catalysts were characterized for their morphology using high resolution scanning electron microscopy (HRSEM) on a Nova nano field-enhanced SEM operating at 5 kV. The SEM samples were prepared by coating them with 5 nm gold layer using a precision etching coating system (Gatan Model 682, Germany). Element distribution within the catalyst sample was determined by energy dispersive X-ray spectroscopy (EDX) built in with the SEM. Nitrogen adsorption experiment was conducted through Brunauer, Emmett and Teller (BET), NOVA®-e Series Model 25 Quantachrome Instruments. The samples were dried at 90˚C for 4 hours in order to desorb the impurities on the surface. The surface area of as-received zeolite particles and calcined zeolite fibers was calculated through multipoint BET method. Data in the relative pressure range of 0.05-0.30 was used for this purpose. Catalyst acid site characterization was obtained using Fourier-transform infrared spectroscopy (FT-IR) on an Agilent Cary 630 spectrometer. Finger prints of dry catalyst samples were collected first. Then, a few drops of pyridine were successively introduced onto the samples before recollecting the spectra. FT-IR spectra were obtained at ambient conditions. 2.3 Feed Preparation HDC experiments were run using de-fined AVR feedstock. Fines were removed in order to avoid potential catalytic role18. Initially, the as-received AVR was heated in a gravity convection oven (model: DX300, Yamato Scientific America Inc., CA, USA) to 150°C in order to reduce its 5 ACS Paragon Plus Environment

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viscosity, then mixed well by a metallic stirring rod in order to homogenize the sample. The homogenized AVR was de-fined by mixing with toluene at a 40:1 v/v ratio, and filtered using 25 µm VWR filter paper. A rotary evaporator (model: Hei-Vap value digital HL/G3, Heidolph instruments GmbH & Co. KG, Germany) was used to evaporate the toluene, and recover the de-fined oil samples, at 93°C and 2 psi. The de-fined AVR was dried for 5 min in the rotary evaporator after complete evaporation of the toluene, and heated to reduce its viscosity and collected as the feedstock to the experiments. The properties of the feed are listed below in Table 118. Table 1: Properties of the de-fined AVR feedstock. °API gravity at 24°C Viscosity at 37°C (cP) Asphaltenes (wt.%) Maltenes (wt.%)

4.4 ± 1.1 >200,000 31.6 ± 3.2 68.4 ± 4.9

2.4 Upgrading The upgrading experiments were operated in an autoclave, i.e. without allowing gases to leave the batch reactor. In a typical run, 50.0 g of the de-fined feedstock, along with the catalyst, whenever applicable, were fed into a 100 mL Parr reactor (1.3 in. i.d. and 4.6 in. length, 4590 Micro Bench Top Reactor, Parr Instrument Company, IL, USA). The zeolite catalysts were mixed with the oil at a concentration of 10,000 ppm. An electrical heating jacket was placed around the reactor, and the apparatus was connected to a control unit. Before any heating began, the reactor was purged with H2 for 3 min, and initially pressurized to 70 or 365 psi, depending on the experiment. For the TC runs, N2 was used to purge the reactor at the beginning with continuous venting and no pressure buildup18, followed by closing the inlet and the exit valves of the reactor, respectively. The pressure inside the reactor was monitored for 5 min to check for leaks, then heating began at a rate of 25°C/min. The temperature, pressure


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and energy consumption were recorded every 5 min. Energy consumption was recorded using a DI-50E meter (Electro-Meters Company Ltd., Alberta, Canada) by connecting it in between the controller and the heating jacket. Mixing at 500 rpm started once the reactor temperature reached 120°C. Reaction time was taken from the point when the reactor reached 300°C. The reaction was allowed to continue for 1 h after reaching this point. At the end of the reaction, heating was stopped, and the reactor was immediately quenched in a water bath to halt any residual cracking reactions. The reactor was then allowed to cool to 24°C, and the final pressure was recorded. Any uncondensed gases were slowly vented from the reactor. The reactor was then weighed and the difference in weight between the initial and the final masses of the reactor was used to calculate the gas yield. More details on the TC experiments can be found elsewhere18. 2.5 Product characterization The liquid product was collected and its viscosity at 37°C was determined using a Brookfield digital viscometer (Model: LVDV-1 PRIME, Brookfield Engineering Laboratories Inc., MA, USA), and its °API gravity at 24°C was measured using a specific gravity bottle (Thomas Scientific, NJ, USA). The product oil was then washed with toluene at a 40:1 v/v ratio of toluene to oil and filtered using 25 µm VWR filter paper in order to reject any toluene insolubles (TI). The filter paper was washed several times with toluene until the filtrate appeared colorless. The filter cake was dried, weighed and collected in order to characterize the TI fraction. Toluene was then separated out of the filtrate using a rotary evaporator at 93°C and 2 psi. During the evaporation process, some volatile components in the product oil were evaporated along with the toluene, and therefore lost from the liquid product. These losses were accounted for through the difference in the mass of the oil sample observed before and after evaporation. The viscosity at 37°C of the remaining de-fined oil product was measured, and then the asphaltenes fraction was rejected by washing with n-heptane at a 40:1 v/v ratio and filtering using 25 µm VWR filter 7 ACS Paragon Plus Environment

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paper. The filter paper was washed several times with n-heptane until the filtrate appeared colorless. The filter cake was dried, weighed and collected in order to characterize the asphaltenes fraction. n-Heptane was removed from the filtrate using the rotary evaporator operating at the conditions described above, and the losses of the volatiles from the maltene fraction were recorded as the difference in the weight of the sample before and after evaporation. The remaining maltene fraction was collected and its viscosity at 37°C and °API gravity at 24°C was measured. Three replicates were prepared for 60% of the runs and the 95% confidence intervals were provided. For selected experiments, the product oil was analyzed using high-temperature simulated distillation (Agilent GC, Alberta, Canada) following ASTM D7169-2005 standard procedure. The conversion of the residue, 545°C+, is of interest in upgrading and was obtained as follows18: =

%,

%, %,

(E1)

3. RESULTS AND DISCUSSION Energy consumed as heat loss from the reactor unit or in the form of sensible heat confirmed that the reactor setup provides reliable estimates of the heat of reactions18. 3.1 Catalyst Characterization Figure 1a,b shows HRSEM images of the as-prepared calcined zeolite-based Ni-W particles and fibers, respectively. Fiber diameter of about 2.0 ± 0.5 µm was observed. Agglomeration can be seen in the particle type catalyst, which is usually the common cause of loss of accessibility to the active sites. From Figure 1c, it is apparent that a uniform elemental distribution is achieved within the fibers14, whereas the particles in Figure 1d do not show such uniform distribution. The silicon, aluminum and oxygen signals are attributed to zeolite while that of


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nickel, tungsten and oxygen are associated with the NiO and WO3 hydrogenation components22. The fibers portray shorter diffusion length between the acid and the metal sites and, in principle, should provide better selectivity toward hydrogenation. The highlighted yellow section of Figure 1d corresponds to areas lacking in metal sites or acidic sites on the particle catalyst. Longer diffusion path for the reacting molecules is, therefore, encountered11. Consequently, poor hydrocracking performance should be expected. Figure 1e confirms that the fiber and the particle catalysts have the same overall metal content. This supports the above conclusion that uniform site distribution, and not merely their overall concentration, is key to better HDC performance. Figure 2 shows the N2 adsorption-desorption curves for the particle and fiber catalysts. The N2 adsorption-desorption isotherms for the fiber and particle morphologies are shown in Figure 2 together with the external surface area and micro pore volume. The isotherms fit closely to IUPAC type H3 hysteresis, which likely suggests non-rigid aggregates23. Such an isotherm does not provide reliable estimate of pore-size. Nevertheless, it is highly likely that the uniform elemental distribution of the fiber morphology contributed to better HDC performance as evident from the higher residue conversion at higher hydrogen pressure Figure 3 shows FT-IR spectra of catalyst with and without absorbed pyridine. The IR spectra of the particle and fiber form catalyst appear similar. Peaks in the 700-800 cm-1 range correspond to asymmetrical stretching within the zeolite structure, as well as benzene substitution when pyridine is present in the samples24. Pore opening vibrations produce the major IR band seen in the 1050-1150 cm1

range25. Acidity of the catalysts can be characterized through the IR bands present at 1455

cm-1 (associated with Lewis acid sites (LAS) due to pyridine absorption), and at 1545 cm-1 (assigned to pyridinium ions formed through proton transfer from Bronsted acid sites (BAS))26. Thus, both fibers and particles possess LAS and BAS, with stronger BAS than LAS due to increased intensity27. Barring other effects such as diffusion limitations, the cracking


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performance of the two forms of catalyst should be similar due to having about the same number of acid sites and strength28. This, again, suggests that the uniform elemental distribution of the fiber morphology contributed to better HDC performance.


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Figure 1: HRSEM images of the as-prepared calcined zeolite-based Ni-W (a) particles, and (b) fiber catalysts. (c) EDX mapping of the fiber catalyst showing uniform elemental distribution. (d) 11 ACS Paragon Plus Environment

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EDX mapping of particle catalyst with yellow circles highlighted areas of non-uniform elemental distribution (e) EDS peaks for the elements detected and the corresponding Ni and W

wt.%.

Figure 2: N2 adsorption desorption curves for zeolite based Ni-W particles and fibers.

a)

b)

Figure 3: FT-IR spectra of catalyst (black) and catalyst with absorbed pyridine (blue), a) particles, b) fibers. 12 ACS Paragon Plus Environment

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3.2 Temperature, pressure and energy profiles The real-time temperature and cumulative energy consumption are given in Figure 4 for all catalytic and control TC and HDC experiments; at low and high initial pressures of H2. The pressure buildup during the reaction is shown in Figure 5. While the temperature and the energy consumption profiles are nearly identical for all of the experiments, the pressure buildup and the final pressure in the reactor after cooling were different. Similar energy consumption for catalytic and control HDC runs suggests that the major endothermic reactions in presence and absence of the catalyst are similar. Moreover, since a similar energy profile is also observed for the TC control run, this suggests that endothermic TC reactions may be the dominant upgrading reactions. Two major reactions typically occur during thermal cracking, 1) endothermic scission reactions and 2) exothermic condensation reactions29. It appears from these results that scission reactions dominated upgrading, even under HDC conditions. Under HDC conditions, an imbalance of the metallic and acidic functions could lead to more emphasis on TC11. However, as seen by the elemental mapping of the fiber catalyst, such an effect is unlikely, although possible for the particle catalyst. Alternatively, insufficient hydrogen pressure may have limited the HDC reactions and the catalyst performance, i.e. mass transfer limited reactions. Typically, batch HDC reactions are run at much higher pressures, on the order of 1,500 psi or greater30. Given the pressure buildup for the different experiments, it is obvious that the catalytic HDC runs promoted hydrogen capping of cracked molecules leading to the formation of gaseous products, as in low pressure HDC or higher chain products, as in high pressure HDC runs31. This precludes mass transfer limitations and confirms the fact that energy consumption was dominated by the TC scission reactions. For the low H2 pressure experiments, during the HDC control experiments, the pressure began to level off after 20 min, despite energy still being consumed at a constant rate. This suggests that the gas formation reactions were limited, perhaps due to the absence of 13 ACS Paragon Plus Environment

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catalyst32. On the other hand, for the TC control experiments, the pressure buildup varied linearly with time between the start and the end of the cracking process. Despite the higher H2 initial pressure of 365 psi, little pressure buildup occurred compared with the low H2 initial pressure experiments, especially toward the end of the run. This could potentially be due to similar rates of hydrogen consumption and gas formation during the cracking experiments, which also explains pressure leveling off in about 10 min from the start of the reaction, especially for the catalytic HDC runs.

500 450 400

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HDC

Figure 4: Mean temperature (solid lines) and cumulative energy (dashed lines) consumption for TC control and lumped HDC catalytic and control experiments. ti is the time when the reactor reached 300°C.


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700

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Figure 5: Pressure buildup during upgrading. ti is the time when the reactor reached 300°C, and tf is the time when the reactor reached room temperature after quenching. Experiments at low H2 pressure (LP) and high H2 pressure (HP) were ran at 70 and 365 psi, respectively. Control (square), zeolite fibers (triangle), zeolite particles (circle). 15 ACS Paragon Plus Environment

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3.3 Product yield

Figure 6 compares the yield of the different upgrading products between catalytic and control HDC, at low (70 psi) and high (365 psi) initial H2 pressures, along with the TC control run. 3.3.1 Control experiments A comparison of the product yields between the TC and low pressure HDC samples reveals major similarities despite the different experimental conditions. While, TC control experiment produced a higher TI yield (4.9 ± 0.2 wt.%) compared to the HDC control (2.3 ± 0.4 wt.%), the low pressure HDC control experiment produced a higher gas and asphaltene yields (3.7 ± 0.3 wt.% and 22.9 ± 0.7 wt.%, versus 1.9 ± 0.4 wt.% and 20.7 ± 0.7 wt.% for the TC control, respectively). The lower TI yield could be attributed to the hydrogen and its role in limiting coke formation. By capping the asphaltenes radicals during cracking, polymerization into coke and its precursors is prevented33. Differences in gas yield seen between Figures 5 and 6 could be due to different components present in the gas. In presence of H2, a higher emphasis on short chain components is expected due to its radical capping function4. Nevertheless, in general the TC control and low pressure HDC control samples provided similar results in terms of product yields, most likely due to the less effective utilization of H2 in the absence of a catalyst12. On the other hand, the high pressure HDC control sample displays more than double the gas yield than of the low pressure HDC run (8.8 wt.%). The TI and asphaltene yields were similar to that of the other control experiments. It appears that the gas yield increased at the expense of the maltenes yield as H2 more effectively capped the short chain radicals. This shift to un-condensable gases is evident in the higher final pressure after cooling the reactor; 230 psi as shown in Figure 5, as well as a reduction in the volatiles yield, i.e. total loss, when compared to the TC or low pressure HDC control samples.


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3.3.2 Catalyst performance at low H2 pressure For H2 pressure of 70 psi, the fiber catalyst showed a relatively high TI yield of 5.5 ± 0.9 wt.%, however slightly lower asphaltenes yield than the particles. This may be attributed to the higher surface area (surface area per BET equation= 310 m2/g; external surface area= 70 m2/g) of the fibers which in turn provides more accessibility to the acidic functional sites relative to the particle-based catalyst (surface area per BET equation= 275 m2/g; external surface area, 40 m2/g), which was more susceptible to agglomeration34. At low H2 pressure both catalysts contributed to higher asphaltenes and lower maltenes yield than that of their control experiments. Although this is an undesirable result, a comparison between the quality of the liquid yield is very relevant before a conclusion could be made on the efficacy of the upgrading process. 3.3.3 Catalyst performance at high H2 pressure At 365 psi, the fiber displayed excellent performance in terms of liquid product yield. While the TI yield (2.3 wt.%) was similar to that of the control run (2.5 wt.%), the asphaltene yield was only 16.9 wt.%, compared with 20.8 wt.% and 22.9 ± 0.7 wt.% for the control HDC high and low pressure runs, respectively. This suggests that the role of the fiber catalyst was more pronounced under higher H2 pressure. Unlike the fiber run at low H2 pressure, hydrogen capped the radicals in the pseudophase more effectively, preventing coke formation. Moreover, many volatiles were produced during cracking leading to total maltenes yield of 75.7 wt.%. The gas yield of 5.2 wt.% was lower than that of the high pressure HDC control runs, as reflected by a final pressure after cooling of only 120 psi, as seen in Figure 5. This result suggests that H2 consumption in the presence of the fiber catalyst was the highest and/or selectivity for the gaseous products was low35. However, the particle catalyst did not perform well at high pressures, providing yields similar to that of the high pressure HDC control.


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0 Gas TC Control

TI HDC Control - LP

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ZF - LP

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Figure 6: The yield (wt.%) of the different upgrading products. Total loss is associated with the loss of volatiles during solvent evaporation. Experiments at low H2 pressure (LP) and H2 high pressure (HP) were ran at 70 and 365 psi, respectively. 18 ACS Paragon Plus Environment

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3.4 Product Quality Figures 7 and 8 show the viscosity and °API gravity measurements, respectively, of the product oil, the combined maltenes and asphaltenes fractions, and the maltenes fraction for all the runs. Figure 7 shows the high temperature simulated distillation results of the product oil, with various distillation cuts shown, namely volatiles (T

Hydrocracking of Athabasca VR Using NiO-WO3 Zeolite-Based

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