Coupling Process of Heavy Oil Millisecond Pyrolysis and Coke

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Coupling Process of Heavy Oil Millisecond Pyrolysis and Coke Gasification: A Fundamental Study Jinhong Zhang, Yuanjun Che, Zhenbo Wang, Yingyun Qiao, and Yuan-yu Tian Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.6b01362 • Publication Date (Web): 26 Jul 2016 Downloaded from http://pubs.acs.org on July 29, 2016

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Coupling Process of Heavy Oil Millisecond Pyrolysis and Coke Gasification: A Fundamental Study Jinhong Zhang, Yuanjun Che, Zhenbo Wang, Yingyun Qiao*, Yuanyu Tian*

State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Qingdao, 266580, China

ABSTRACT

To realize the clean and high-efficiency utilization of the inferior heavy oil, a novel conceptional process coupling heavy oil millisecond pyrolysis and coke gasification processes was proposed. The effects of reaction temperature and residence time on the flash pyrolysis behavior of vacuum residue (VR) were firstly investigated using Py-GC/TOF-MS. Results show that increasing the reaction temperature from 600 to 800 °C, as well as shortening the residence time from 400 to 200ms could reduce the coke yield by approximately 28%. 1-alkenes are the most abundant class in the cracked products, and the selectivity of total olefins in the light products from C3 to C16 could be over 75%. Then the cracking performance of VR over Y zeolite based equilibrium FCC catalysts, silica sands, and calcium aluminate catalysts were studied by a fluidized bed reactor. Furthermore, the coked catalysts were gasified in steam and steam-O2 mixture atmosphere, respectively. Results show that the calcium aluminate catalyst has appropriate cracking activity, excellent light olefin selectivity, significant coke gasification activity and good hydrothermal stability. The feed conversion can be up to approximately 98 wt % at 650 °C, with 70% light olefin selectivity in the cracking gas, while the coke yield is only 6.4 wt % in the pyrolysis of VR 1

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with a 16 wt % Conradson carbon residue. Moreover, the calcium aluminate catalyst could reduce the gasification temperature and increase the coke gasification efficiency and the hydrogen content of syngas. These studies demonstrate that directly processing inferior heavy oil in the coupling process using calcium aluminate catalyst is feasible.

KEYWORDS

Heavy oil; Millisecond pyrolysis; Coke gasification; Fluid catalytic cracking (FCC); Basic catalysis

1. INTRODUCTION

Heavy oil resources, including oil sand, oil shale, and extra-heavy oil, account for a large part of the world’s petroleum resources. In view of the progressively increasing heavier nature of the crudes, the quantity of heavy residues is expected to increase in the future.1 As heavy crudes are commonly cheaper than light crudes, a trend of processing heavier crudes is catching up fast in the refineries.2 Fluid catalytic cracking (FCC) units play an important role in heavy oil upgrading, which account for approximately 24% of worldwide commercial residue processing capacity.3 However, heavy oil commonly has the characteristics of high boiling point, high contents of resins and asphaltenes, high contents of heteroatoms (i.e., Ni, V, S, N, O), and high coking tendency. It will create huge challenges and opportunities in the heavy oil processing technology.

As the heavier nature of feedstocks, the increasingly strict environmental regulations, and the demand for high value petroleum products, FCC technologies are mainly developing 2


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under the following six directions: (1) increasing the processing capacity of inferior feedstocks; (2) reducing the dry gas and coke yields, meanwhile increasing the light oil yield; (3) improving the product quality; (4) producing more petrochemical feedstocks, such as ethylene and propylene; (5) reducing the SOx and NOx emissions during the generation of coked catalyst; (6) rational utilization of coke.

The major limitation of RFCC process is the need of good quality feedstocks, which contain relatively low amounts of metals, sulfur, nitrogen and asphaltenes.4 Even though many processes5-7 and catalysts8 have been developed, the blending ratio of inferior feedstocks is still limited due to the high coke yield and catalyst deactivation.9 Recently, Zhang et al.10,

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studied a residue cracking gasification process, which used spent FCC

catalyst and kaolin to pretreat vacuum residue (VR), and steam gasification of the coke to produce syngas. However, it should be noted that due to the poor hydrothermal stability, the typical FCC catalyst cannot be operated in commercial units. Thus, novel poisoning-resistant catalysts with higher hydrothermal stability need to be developed.

In the typical RFCC process, the gas oils mixed with residues and cycle oil are cracked in one riser reactor in approximately 3 seconds. Shan et al.12-14 proposed a two-stage riser (TSR) FCC technology which divided one riser reactor into two shorter riser reactors and shortened the residence time to 1-2s. As the cracking gas can be separated and the catalyst can be regenerated in time, the increased feed conversion and light oil yield, decreased dry gas and coke yields, as well as improved product quality can be achieved. However, due to the back-mixing of oil gas and catalyst, it is difficult to realize the very short contact time in a

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riser reactor. Downer reactor, in which gas and solids flow downward concurrent, is closer to the plug flow reactor compared to the riser reactor. Therefore, it is more suitable for the high severity operating processes, such as the high temperature, ultra-short reaction time with the intermediates as the desired products.15 Pilot test results showed that at the same conversion, downer reactor obtained higher yields of gasoline and light olefins, lower yields of dry gas and coke.16

Commonly the operating temperature of RFCC process is around 500 °C, the high boiling point residues cannot vaporize and eventually deposited on the surface of the catalyst, leading to the increase of coke generation and catalyst deactivation. Wu et al.17 compared the thermal cracking of n-hexadecane under liquid-phase and gas-phase conditions, and found that gas-phase cracking produced more gas products and no addition compounds, while liquid-phase cracking gave a higher selectivity of addition compounds. Gray et al.18 investigated the thermal cracking of vacuum residue in a continuous flow aerosol reactor, and found that high-temperature (700-800°C) and short-residence time (100-115ms) was conductive to reducing coke yield, as more fraction of feed that would vaporize and react in the vapor phase. Therefore, it is expected to realize the high-temperature millisecond pyrolysis of heavy oil in the downer reactor.

In the conventional FCC process, a certain amount of feed sulfur and nitrogen would deposit on the coke and eventually be oxidized into SOx and NOx in the regenerator. Research found that the higher concentrations of sulfur and nitrogen in the feed would increase the percentage of feed sulfur and nitrogen existed in the coke.19, 20 Thus, when processing inferior

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feedstocks with higher sulfur and nitrogen contents, the FCC regenerator poses a challenging to the increasingly strict environmental regulations for controlling SOx and NOx emissions.

Compared with just burning in the regenerator, coke gasification is a more rational way for coke utilization, which cannot only produce the hydrogen-rich syngas as byproduct, but also reduce the emission of pollutants. As during the gasification process, most of the sulfur and nitrogen in the coke are converted to H2S and N2, respectively.21 Due to the low reactivity of coke, taking proper gasification agent and catalyst to increase the carbon conversion is necessary. Tian et al.22 found that gasification with steam-O2 mixture could significantly reduce the apparent activation energy. Corma et al.23 found that simultaneous steam reforming and combustion could enhance the coke conversion, and hydrotalcite had a much higher steam reforming activity than typical FCC catalyst. Pant et al.24 used calcium aluminates as the catalyst for steam pyrolysis of n-heptane and found that the conversion and light olefin yields were significantly increased compared to non-catalytic pyrolysis. However, the heavy oil cracking activity and coke steam/steam-O2 gasification performance on calcium aluminate catalysts have not been studied yet.

In this work, a heavy oil millisecond pyrolysis and gasification (MSPG) process was proposed to realize the high-efficiency and clean utilization of inferior heavy oil. As shown in Figure 1, the heavy oil is injected from the top of a downer-type pyrolysis reactor and contact with the high-temperature regenerated catalyst. After cracking in milliseconds, the oil gas and coked catalyst are separated by a horizontal cyclone separator. Then the oil gas is piped to a distillation system and cut into dry gas, liquefied petroleum gas (LPG), gasoline, diesel, and

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heavy cycle oil (HCO). Light olefins can be further obtained from dry gas and LPG. After stripped by steam, the coked catalyst is regenerated by steam and air mixture in a riser-type gasifier. The regenerated catalyst is recycled to the downer reactor, while the syngas can be further processed by a water-gas shift reactor to increase the H2 yield. After removed off CO2, the remained H2 can be further used to upgrade gasoline and diesel to improve the product quality. This process could reduce the coke yield, increase the selectivity of light olefins, produce H2-rich syngas and, furthermore, reduce the emission of pollutants.

The purpose of this paper is to understand the fundamentals of the heavy oil millisecond pyrolysis and the gasification of generated coke. In the first work, the effects of residence time and temperature on minimizing the coke yield during vacuum residue millisecond pyrolysis were investigated using an analytical Py-GC/TOF-MS. Then the cracking performance of VR over typical FCC catalyst, silica sand and calcium aluminate catalyst was studied by a fluidized bed reactor. Finally, the influence of catalyst on coke conversion and syngas composition was evaluated by TG and fluidized bed gasification.

2. EXPERIMENTAL SECTION

2.1. Feedstock and Catalyst.

A vacuum residue (VR) provided by Shengli oil field (Shandong Province, China) was used as the heavy feedstock and its properties are shown in Table 1. The VR has high density and viscosity. Moreover, the Conradson carbon residue (CCR) of the VR is up to approximately 16 wt %.

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The catalysts used in the experiments were commercial Y zeolite based equilibrium FCC catalysts, silica sands, and commercial calcium aluminate catalysts. The properties of these catalysts were shown in Table 2.

2.2. Experimental Apparatus and Product Analysis.

2.2.1. Py-GC/TOF-MS

The Py-GC/TOF-MS experiments were conducted using a Pyroprobe 5200 Series pyrolyser (CDS Analytical Inc.) with direct connection to a DANI gas chromatograph coupled to a time of flight mass spectrometer. Each experiment was conducted with 0.400±0.005 mg of oil sample in a quartz boat. An analytical balance with the readability of 0.001 mg was used for weight, to ensure the accurate quantity of samples and residues. In the Py-GC/TOF-MS apparatus, the residence time refers to the time of gas going through the pyrolyser, which is controlled by the flow rate of carrier gas, while the reaction time refers to the time of pyrolyser remained at the final temperature, which can be set from 0-60 s by the control software.

The pyrolysis was carried out at the set temperature from 600 °C to 800 °C with a heating rate of 2000 °C/s. The GC separation of pyrolysis vapors were done by a DB-5MS column (60m × 0.25mm, 0.25µm film thickness) with Helium (99.999%) carrier gas (1mL/min). The GC inlet was 280 °C and a split ratio of 50:1 was used. The oven was programmed to start at 30 °C (hold for 15 min) and then ramp to 60 °C with 1 °C/min (hold for 1min), and then raise to 200 °C with the heating rate of 3 °C/min followed by 10 °C/min up to a final temperature of 280 °C. The mass spectrometer was operated in EI mode at 70eV, and the mass spectra 7


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were obtained from m/z 35 to 500. The chromatographic peaks were identified according to the NIST mass spectral library.

Due to the very complex components in this study and the lack of commercially available standards, the Py-GC/TOF-MS technique could not give the direct quantitative analysis of the pyrolytic products. However, the chromatographic peak area of a compound is considered linear with its quantity. Thus, the peak area value of each product can be compared to reveal the changing of its yields under different reaction conditions.

2.2.2. Fluidized bed experimental apparatus

The schematic diagram of fluidized bed experimental apparatus is shown in Figure 2, which is comprised of an oil feeding system, a reaction system, a gas supplying system, a product collecting system, and a product analyzing system. In the VR cracking tests, about 240 g catalysts were first loaded in the reactor and then heated to the desired temperatures. Meanwhile, VR and steam was preheated to about 200 °C and 350 °C, respectively. For each test, the VR was fed at a rate of 20 g/min for 1.5 min, while that for water was 10 g/min. The oil and steam were mixed in a furnace and then continuously fed into the reactor by pumps. The fluidized bed reactor was made of stainless steel and its total length was 700 mm with a 350 mm long expanded section of 102 mm in diameter. The oil-steam mixture was atomized into tiny oil droplets by a nozzle and then injected into the bottom of the reactor. When the oil droplets flowed upward to the catalyst bed, the heavy hydrocarbons were cracked into light fractions over the catalyst. Then the formed oil gas and coked catalyst were separated by a steel filter in the expanded section of the reactor. The cracking gas was cooled by a two-step 8


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condensation process. The collected liquid products were weighed to determine the liquid yield then analyzed by a gas chromatograph (GC) for the fraction distribution, while the non-condensable gas was measured by a gas flow meter then analyzed by another GC for the gas composition. Meanwhile, the coke yield can be obtained by measuring the carbon content of the coked catalyst. After the cracking process, the coked catalysts were stripped by steam for half an hour and then gasified in the fluidized bed reactor by steam or steam-oxygen mixture. For each test, about 100 g of coked catalysts were loaded in the reactor. The nitrogen was first used to fluidize the catalyst until the temperature reached to the preset value, and then the gas injection was switched to the pure steam (4ml/min) or steam-oxygen mixture (98% steam + 2% O2). The generated gas was measured by a gas flow meter then analyzed by GC. In this study, each test was repeated at least twice, the mass balance was over 95% and the relative error of the measurement was less than 5%.

2.2.3. Analysis and Characterization

The compositions of cracking gas and syngas were analyzed by a Varian CP-3800 gas chromatograph (GC), while the liquid products analyzed for the simulated distillation by another Varian CP-3800 GC according to the ASTM-D2887 standard test method. The cut points were set at 204 °C for gasoline, 350 °C for diesel, 500 °C for vacuum gas oil (VGO), and higher than 500 °C for heavy oil. The conversion was defined as the yield sum of dry gas, liquefied petroleum gas (LPG), gasoline, diesel, VGO and coke. As the feedstock is vacuum residue, VGO is considered as the product. To determine the coke yield, the carbon deposited on the catalyst was measured with a coke analyzer, which burned coked catalyst samples in

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pure oxygen and then analyzed the CO2 content by a GC.

Nitrogen adsorption and desorption isotherms were performed at 77 K on a volumetric adsorption system (Micromeritics ASAP 2020, American) to determine the specific surface area and pore volume. The X-ray diffraction characterization was carried out on an X’Pert MPD (DANalytical Co.) analyzer using Cu Kα radiation at 40 kV and 40 mA, with a scanning speed of 10o/min. Thermogravimetric (TG) analysis of coked catalysts were performed using a Netzsch STA 449F3 thermal analyzer system by using about 200 mg of sample in each run at heating rates of 10 °C/min in a range from ambient temperature to around 1100 °C under steam atmosphere (40 ml/min). As the low coke yield of samples, more samples need to be used in this experiment. To eliminate the negative effect on heat and mass diffusion caused by large sample weight, a larger pot (DN. 18mm) was used to make samples more uniform distribution with thinner layer.

3. RESULTS AND DISCUSSION

3.1. Effect of Residence Time on the Millisecond Pyrolysis of Heavy Oil.

Research had found that a shorter residence time, in the range of 0.6 to 3.2 s, is beneficial to reduce the coke yield in the FCC process.25 However, it has not yet been studied in the millisecond level. In this work, the Py-GC/TOF-MS apparatus was used to investigate the millisecond pyrolysis behavior of Shengli VR.

In the Py-GC/TOF-MS apparatus, the residence time is controlled by the flow rate of carrier gas. According to the descriptions, the flow rate of carrier gas can be varied from 10 to 10


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30 mL/min. The flow rate of 15, 20 and 30 mL/min is corresponding to approximately 400, 300 and 200 ms, respectively. To ensure heavy oil is fully converted, the pyrolyser is kept at final temperature for 10 s. Figure 3 shows the effect of residence time on the coke yield at 600, 700 and 800 °C. When the residence time increased from 200 to 400 ms, the coke yield all showed upward trends. It also can be seen that the increases of coke yield up to approximately 16% at 600 °C, which is twice of that at 700 and 800 °C. This is because of the higher temperatures enhanced the gas-phase reactions which is beneficial to reduce the formation of addition compounds.17 Moreover, a higher temperature could enhance the evaporation of heavy fractions, leading to a reduction of film thickness. Thus, more cracked products can be transported out of the liquid phase, instead of forming coke.

This work showed that a shorter residence time, even in millisecond level is also helpful to inhibit the generation of coke, which would be taken into account in the designing of pilot and industrial units. Moreover, the reaction temperature has a significant effect on the coke formation. Increasing the reaction temperature from 600 to 800 °C, as well as shortening the residence time from 400 to 200ms could reduce the coke yield by approximately 28%.

3.2. Effect of Reaction Temperature on the Millisecond Pyrolysis of Heavy Oil.

It is well known, reaction temperature is an important parameter for heavy oil cracking, which has significant influence on the reaction rate, product distribution and product quality. In this work, the effect of reaction temperature on VR pyrolysis process under millisecond level was investigated by Py-GC/TOF-MS.

3.2.1. Residue Yield. 11


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In the heavy oil pyrolysis process, cracking and condensation are two key reactions. The cracking reaction is endothermic and has lower activation energy, while the condensation reaction is exothermic and has higher activation energy. On the one hand, a relative high temperature is beneficial to the condensation reaction according to the mechanism; on the other hand, it could also enhance the gas-phase cracking to inhibit the generation of coke as more high-boiling point fractions could be vaporized. Moreover, due to the increased vaporization rate, the film thickness could be reduced. Dutta et al.26 found that as the feed film thickness decreased from 150 to 15 µm, the coke yield decreased from 23 to 18 wt % at 510-530 °C. Furthermore, the reaction rates of condensation reactions are generally lower than that of cracking. Therefore, if the high-temperature reaction could be controlled in a short-residence time, the condensation reaction could be inhibited, leading to a decrease of the coke yield.

Figure 4 shows the variation of residue yield with the reaction time of 600, 700 and 800 °C. As the increase of reaction time, the residue yields all decreased sharply, and then became stable. But, the reducing rates and final values varied with the temperature. When the reaction temperature increased from 600 to 800 °C, the final residue yield decreased from 10.7 to 9.2 wt %, which indicates that approximately 14% of residues were further cracked, and the high-temperature millisecond operation is beneficial to reduce the coke yield. This is different from that in the conventional FCC process, where the coke yield increased monotonous with the temperature.14,

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This may be due to the inhibition of the second

condensation reaction in the millisecond pyrolysis process.

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According to Arrhenius equation, the raise of reaction temperature will lead to the increase of reaction rate. As the decrease of reactant concentration is basically proportional to the mass loss, the ratio of reaction rate within the same reaction time at different reaction temperatures can be defined as follow: , ,

=

,

,

(1)

The calculated results are shown in Figure 5. It can be seen that at the initial time a higher temperature significantly enhanced the cracking reaction of VR. The reaction rate at 700 °C is 3.5 times faster than that of 600 °C in the first 1s, and the value at 800 °C is 2.4 times faster than that of 700 °C in the first 0.5s. With the reaction progress, the reactant concentration decreased, the reaction is controlled by the reactant concentration. Thus, the effect of temperature is not obvious when the reaction time longer than 3s. However, to avoid the overcracking of desired products, even in the conventional FCC process the reaction time is shorter than 3s. Therefore, the millisecond pyrolysis should be operated at a higher temperature to make sure a relative high feed conversion.

3.2.2. Pyrolysis Products.

Figure 6 shows the ion chromatograms from millisecond pyrolysis of Shengli vacuum residue (VR) at 600, 700 and 800 °C, respectively. There are more than 300 peaks can be identified by GC/TOF-MS analysis. Table 3 lists some typical components of pyrolysis products according to the carbon numbers from 3 to 16. The lack of methane, ethane, ethylene and propane is due to the fact that the analysis apparatus cannot identify compounds below C3. One drawback of the Py-GC/MS apparatus is that it did not allow the product 13


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collection, thus the exact products yield could not be determined. However, because of the chromatographic peak area for a compound is in proportion to its yield, it is able to have a primary estimate of the yield changes through the comparison of the chromatographic peak areas obtained at different pyrolysis temperatures.28 A complete listing of the selectivity of identified product groups is presented in Table 4. The values are normalized on the recognized fraction of the total area of the chromatogram. As the 1-butene/isobutene and propene/cyclopropane/propane cannot be separated very well, butene and propene were list separately.

Figure 7 shows the main reaction pathways for the millisecond pyrolysis of VR. The crackable hydrocarbons in VR (such as alkanes, alkyl-chains and aromatics with 1-3 rings) are mainly cracked into alkenes and small amount of alkanes and aromatics, while most of the polyaromatics finally form coke. The formed alkenes can be further cracked into alkadienes or generate aromatics undergo aromatization reactions at high temperatures. Also, alkadienes can form aromatics through aromatization at high temperatures.

Alkenes. Alkenes are in each case the most abundant classes in pyrolysis products. It can be found that in the millisecond pyrolysis process, 1-alkenes are the major products, which is conforming to the free radical theory.29 Figure 8a shows that the peak areas of alkenes were increased monotonically as the increase of pyrolysis temperature from 600°C to 800°C, which indicates a higher temperature is beneficial to the generation of alkenes. The peak areas of olefins were all increased sharply, and then decreased gradually with the growth of the carbon number. The maximum values were all obtained at C6 for different temperatures. 14


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When the pyrolysis temperature increased from 600 °C to 700 °C, the peak area of olefins from C3 to C16 all showed a significant rising, especially for C5 and C6; however, as the temperature further went up to 800 °C, the increase became slower. Therefore, the selectivity of alkenes reached the peak value at 700 °C, which can be seen from Table 4. This is because a high temperature enhanced the further cracking and aromatization reactions of alkenes to form alkadienes and aromatics.

Alkanes. It can be seen from Table 4, the selectivity of alkanes significantly decreased with the increase of reaction temperature. The detailed distributions of alkanes were shown in Figure 8b according to the carbon number. There exist two peaks for all three temperatures. The maximum values were all obtained at C5, which are higher than the second peak values at C13. When the pyrolysis temperature increased from 600 °C to 700 °C, the peak areas of C5-C16 alkanes were all increased significantly, due to the further cracking of high-boiling point fractions. However, as the further increase of the pyrolysis temperature, the cracking reactions of alkanes higher than C7 were also enhanced, which can be inferred from the increasing of peak areas of C4-C7 alkanes and the decreasing of peak areas of alkanes over C8. Thus, if the pyrolysis temperature higher than700 °C, it would lead to the overcracking of gasoline and diesel fractions.

Alkadienes. The data presented in Table 4 show that the content of alkadienes in pyrolysis products significantly increased from 2% at 600 °C to approximately 16% at 800 °C. Figure 8c presents the distribution of diolefins according to the carbon number. For all temperatures, the most abundant compounds are 1,3-pentadiene, 1,3-cyclopentadiene, and their isomers. As

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the increase of reaction temperature, the significantly increased alkadienes abundance is mainly because of the generation of C5 and C6 diolefins from the dehydrogenation of C5 and C6 olefins or the cracking of larger olefins. It is different from the study of Calemma et al.30, the terminal diolefins from C5 to C16 were all detected from 600 to 800 °C, which may be due to the development of TOF-MS.

Aromatics. The data presented in Table 4 show that the content of aromatic compounds in pyrolysis products significantly increased from 1% to 10% with the increase of reaction temperature. Figure 8d shows that the most abundant class is mononuclear aromatics from C6 to C9, including benzene, toluene, styrene, xylene, indene, and mono- and multi-substituted alkylbenzenes with alkyl groups up to C3. One of the reasons for the boost of aromatic compounds is the higher temperature could enhance the pyrolysis of substituted aromatics. The reaction pathways of alkyl-aromatics have been described in the literature.31,

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In

addition to the breakdown of aliphatic bridges, the presence of much intermediates (such as 1,3-Cyclohexadiene, 1-Methyl-1,4-cyclohexadiene, 1,3-Dimethyl-1-cyclohexene, et al.) in the pyrolysis products indicates that the increased aromatics in a large part are due to the cyclization of olefins followed by dehydrogenation reactions.

Alkene/Alkane. As the increase of pyrolysis temperature, the selectivity of alkanes decreased significantly, while that for alkenes reached the peak value at 700 °C. The variations of peak area ratios of alkenes to alkanes with different carbon numbers were shown in Figure 9. The ratios for C5-C16 hydrocarbons monotonically increased from 1.1-2.8 at 600 °C to 2.0-8.8 at 800 °C, while that for C4 hydrocarbons showed the maximum value at 16


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700 °C. This is result from the higher generation rate of alkenes and the overcracking of alkanes larger than C8, which can be seen in Figure 8b. Moreover, the yield of butenes was 18-48 times higher than that of butanes, which indicates the high-temperature and short-residence time operation favors production of light olefins over alkanes. This may be due to a higher temperature enhanced the gas-phase cracking over liquid-phase cracking. Research found that hydrogen abstraction is predominant and much faster in liquid-phase cracking due to a higher reactant concentration. This will enhance the radical addition in the propagation step, leading to the reduction of olefins in the cracked products and the generation of high molecular weight compounds.17, 33-35

Due to the high boiling of vacuum residue, the liquid-phase cracking is inevitable in the pyrolysis process of VR. Therefore, how to fast remove olefins from liquid phase became the key matter for raising the olefin yield and reducing the coke yield. The first method is to reduce the initial film thickness to enhance the mass transfer, which has been investigated by Gray et al.26, 36 Shortening the residence time is another method, which was applied in this work and some previous work.18, 37 However, in the catalytic cracking process, the hydrogen transfer activity of catalyst is also need to be considered.

3.3. Effect of Catalyst on the Pyrolysis of Heavy Oil.

The Y zeolite based catalyst has been commercially used in FCC process for decades due to the high cracking activity, high product selectivity, and relative high hydrothermal stability. The hydrogen transfer activity plays an important role in reducing the olefin content of FCC gasoline,38, 39 and in the catalytic cracking of polycyclic aromatic hydrocarbons.40 However, it 17


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would also lead to the saturation of light olefins and the generation of coke. To produce more light olefins, ZSM-5 zeolites are commonly used because of the shape selectivity and reduced hydrogen transfer activity.41 However, they are difficult to be used to process inferior vacuum residues due to the fast deactivation and low feed conversion. Moreover, conventional FCC acid catalysts cannot suffer the steam atmosphere in the coke gasification process.

To inhibit the hydrogen transfer reaction, three methods were used in this work. Because the hydrogen transfer reaction is exothermic, and has a slower reaction rate; it will benefit from a relatively lower reaction temperature42 and longer residence time.43 Therefore, the first method is choosing a relative high reaction temperature. Considering that prolonging the residence time would cause the undesired secondary reactions, especially under higher reaction temperature, a relatively short residence time should be chosen, which the second method is. The high temperature millisecond pyrolysis of VR has been studied by the Py-GC/TOF-MS apparatus. The third method is to reduce the hydrogen transfer activity of catalysts. In this work, a kind of solid base catalyst, calcium aluminate, was tried to directly cracking vacuum residue.

To understand the effect of different catalysts on the conversion of vacuum residue, a kind of commercial Y-zeolite based equilibrium FCC catalyst and pure silica sand were tested as the compared experiments. All the tests were carried out in a fixed fluidized bed experimental apparatus. The operating conditions and product distributions were list in Table 5. According to the Py-GC/MS studies, the reaction temperature should be higher than 600 °C to enhance the gas-phase cracking, but be lower than 700 °C to avoid the overcracking of desired

18


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products. The catalyst-to-oil weight ratio was set as about 8, which is close to the industrial operation condition. To obtain a better atomization effect in the experimental apparatus, the steam-to-oil weight ratio was set as 0.5.

The VR was almost completely converted over Y-zeolite based FCC catalyst at the tested conditions. However, due to the strong acidity, the overcracking of gasoline and diesel fractions also significant, only 46.87 wt % of liquid yield can be obtained. Hence, the Y-zeolite based FCC catalyst is unsuitable for the high-temperature pyrolysis of VR. Catalytic cracking is the major way to convert heavy oil in the conventional FCC process, but as the reaction temperature was as high as 600 °C, the influence of protolytic cracking and thermal cracking could not be neglected,14 which lead to a high yield of methane (about 7%). The high hydrogen transfer activity of the Y-zeolite enhanced the saturation of olefins, leading to a lower olefinicity (defined as the weight ratio of ethylene, propylene and butenes in the cracking gas) of cracking gas. As the low catalytic activity, the reaction temperature over silica sand was increased to 650 °C. Even though the liquid yield over silica sand reached 62%, the desired products, gasoline and diesel, only account for about 37%. The high yields of VGO and heavy oil indicates that non-catalytic thermal cracking is insufficient to convert VR.

The conversion of VR increased significantly in the presence of calcium aluminate catalyst, compared to the silica sand. Lemonidou and Vasalos et al.44 postulated that the cracking reactions of hydrocarbons over calcium aluminate catalyst proceed via free radical mechanism, and the catalytic activity is attributed to the presence of peroxidic active oxygen

19


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which can react with hydrogen at higher temperature, promoting the initiation reactions by abstracting hydrogen from the parent hydrocarbons. Thus, a higher feed conversion and cracking gas yield can be obtained. The olefinicity of cracking gas increased to about 70% with 30.8 wt % total yield of light olefins at 650 °C, while the coke yield decreased to around 6.4 wt %. The coke/CCR ratio is only 0.4, which is much lower than that in the delayed coking process (1.4-1.7). This effect can be described from two aspects. Firstly, it can be attributed to the activation of C-H bonds on the [O](S) surface sites of the catalyst, enhancing the formation of primary radicals ·

[O](S) + CnH2n+2

[OH](S) + CnH 2n+1.

Then the alkyl radical can rapidly react with another [O](S) surface sites to form an olefin molecule45, 46 ·

[O](S) + CnH 2n+1

[OH](S) + CnH2n.

Secondly, as there are no acid sites on the calcium aluminate, the hydrogen transfer reaction can be inhibited. Thus, the calcium aluminate catalyst is beneficial to enhance the conversion of heavy oil and the generation of light olefins, as well as inhibit the generation of coke.

3.4. Coke Gasification.

The major problem for coke combustion is the emission of SOx , NOx and CO2. By contrast, coke steam reforming is a more attractive way, as the low-value coke can be transformed into high-value hydrogen in the process. To enhance the conversion of coke and keep the heat balance, the mixture of steam with oxygen was used as the gasification reagent, which can 20


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significantly reduce the apparent activation energy.

22, 23

Zhang et al.10 found that the desired

coke gasification temperature was about 800 °C when pure steam or a mixture of 95 vol % steam plus 5 vol % oxygen was used. The main reactions occurring in the steam−O2 gasification of coke can be found in the literature.22, 23 Reactions between C/CO and H2O are the desired reactions in this study, which could generate H2. Research found that the two reactions could be accelerated by active intermediates from the metal catalyst, such as K, Na, Ca, Mg, and so on.47-49 Thus, it can be predicted that the calcium aluminate catalyst could enhance the conversion of coke and the generation of H2.

3.4.1. TG analysis.

The effect of catalyst on the gasification of coke was firstly investigated by the TG analysis. Figure 10 shows the TG and DTG curves with the heating rate of 10 °C/min under steam atmosphere. Before the gasification of coke, the weight of samples began to decrease since 400 °C, which may be due to the evaporation of residual oil on the catalysts. The DTG curves illustrate that the gasification of coke on calcium aluminate catalyst started at about 680 °C and reached the peak value at 797 °C, while the coke on silica sand began to gasified at around 720 °C and reached the maximum value at about 882 °C. The DTG value indicates that the maximum coke conversion rate over calcium aluminate catalyst is 1.6 times faster than that over silica sand. It can be seen from the TG curves, the coke on calcium aluminate catalyst was completely converted at about 850 °C, while on silica sand the temperature was as high as 990 °C. The gasification time for coked calcium aluminate catalyst was only 17 min. Thus, the calcium aluminate catalyst could not only reduce the reaction temperature, but

21


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also increase the gasification efficiency of coke.

3.4.2. Fluidized bed gasification.

To obtain the composition of syngas, the coked catalysts were further gasified in a fluidized bed reactor. Table 5 presents the composition of syngas from the coke gasification tests over the silica sand and calcium aluminate catalyst. The coke was gasified at 780-800 °C for about 60 min, using the steam or steam-oxygen (2 vol % O2) mixture as the gasifying agent. As shown in Table 5, approximately 90% of the coke on calcium aluminate could be gasified by steam at 800 °C, which is 7% higher than that at 780 °C, while that for silica sand only reached about 50%. The adding of 2 vol % O2 in steam could further increase the carbon conversion by 9%. The enhancing effect of calcium aluminate catalyst is in accordance with the results of TG analysis. However, in the TG analysis the coke can be completely removed off from the calcium aluminate catalyst by pure steam in 20 min. This may be due to the turbulence in fluidized bed reducing the contacting efficiency of coke with steam, compared to that in the TG analysis.

For silica sand, the H2 account for 54 vol % in the produced gas, while for calcium aluminate catalyst, the H2 content can be further increased to about 59 vol %. This is because of the catalytic effect of Ca, in the form of active intermediates includes CaO and CaO2,50 enhanced the water-gas shift (WGS) reactions to boost H2 production. When introducing oxygen into the reaction, the H2 content significantly reduced due to the partial combustion. Therefore, it is better to choose a more suitable catalyst and optimize the gasification reactor to enhance the coke-steam contacting efficiency, instead of introducing O2 in to the reactions. 22


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Figure 11 shows the XRD spectra of the fresh, aged and regenerated calcium aluminate catalysts. The aged catalyst was treated with steam at 800 °C for 8 h to simulate hydrothermal deactivation, while the regenerated catalyst suffered the VR cracking and steam gasification for 3 times. In the regeneration process, the coked catalyst was gasified by steam at 800 °C for 1 h, and then combusted in oxygen to removing the remaining coke on the catalyst. As shown in Figure 11, the crystal peaks of the three catalysts were almost the same, which indicates that the calcium aluminate catalyst has good hydrothermal stability.

4. CONCLUSIONS

A conceptional process coupling the heavy oil millisecond pyrolysis process and the coke gasification process (MSPG) was proposed in this study. The results show that shortening the residence time in millisecond level is still helpful to inhibit the generation of coke. Appropriately increase the reaction temperature could increase the reaction rate and conversion efficiency significantly. High-temperature millisecond operation for the pyrolysis of VR is helpful to enhance the gas-phase cracking and remove olefins from the liquid phase in time, leading to lower coke yield and higher olefin yield.

An acceptable conversion of VR can be obtained over the calcium aluminate catalyst, which could not only enhance the generation of light olefins, but also inhibit the generation of coke. Approximately 98 wt % of VR were converted at 650 °C, with 30.8 wt % yield of light olefins, while the coke yield was 6.43 wt %. The utilization of calcium aluminate catalyst in the coke gasification process could significantly enhance the conversion of coke, and the H2 content in the syngas can be up to approximately 59 vol %. The calcium aluminate material is 23


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expected to be a desired catalyst for the heavy oil millisecond pyrolysis and gasification process, which could realize the high-efficiency and clean utilization of inferior heavy oil, without worrying about the high heteroatom content, the high coking tendency, and the high SOx and NOx emissions in the regenerator. Moreover, the H2 byproduct could help to secure H2 supplies in refineries.

AUTHOR INFORMATION

Corresponding Author

* Tel.:+86-532-86981711; fax: + 86-532-86981711; e-mail: [email protected] (Y. Tian), [email protected] (Y. Qiao).

Notes

The authors declare no competing financial interest.

ACKNOWLEDGMENTS

The authors acknowledge the financial support provided by the National Natural Science Foundation of China (21576293 and U1462205), the Natural Science Foundation of Shandong

Province

(2014M560589),

(BS2015NJ006), and

the

China

Postdoctoral

Science

Foundation

the Fundamental Research Funds for the Central Universities

(15CX02020A and 15CX05044A).

REFERENCES

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Tables Table 1. Properties of Shengli Vacuum Residue. Table 2. Properties of catalysts. Table 3. Typical pyrolysis products of Shengli VR by Py-GC/MS. Table 4. Percentage hydrocarbon distribution for pyrolysis products. Table 5. Product distribution of VR cracking over different catalysts. Table 6. Gas composition of coke gasification over different catalysts.

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Table 1. Properties of Shengli Vacuum Residue

a

Density (20oC), kg/m3

Viscosity (100oC), mm2·s-1

CCRa, wt %

C, wt %

H, wt %

S, wt %

N, wt %

O, wt %

H/C

980

900

15.70

87.00

12.00

0.26

0.45

0.22

1.66

CCR is the Conradson carbon residue.

32


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Table 2. Properties of Catalysts. Items

Y zeolite based catalyst 3

Packing density (kg/m ) 892 Particle size (µm) 75-150 Surface area (m2/g) 237 3 Pore volume (cm /g) 0.15 Average pore diameter (nm) 4.85 XRF analysis of the catalysts (wt %) Al2O3 47.09 SiO2 42.51 CaO 0.88 MgO 0.72

Silica sand

Calcium aluminate

1282 75-150 -

1399 75-150 10.6 0.06 9.16

99.5 -

48.60 6.92 31.44 0.51

33


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Table 3. Typical pyrolysis products of Shengli VR by Py-GC/MS. No.

Name of compound

Formula

Molecular weight (g/mol)

type

1

Propene

C3H6

42

Alkene

2

1-Butene + 2-Butene

C4H8

56

3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42

1-Pentene 1,3-Pentadiene 1,3-Cyclopentadiene 4-Methyl-1-Pentene 1-Hexene Hexane 1,3-Cyclohexadiene Benzene 1-Heptene Heptane Toluene 2-Methyl-1-heptene 1-Octene Octane 1,3-Dimethyl-1-cyclohexene 1,2,4,4-Tetramethylcyclopentene p-Xylene Styrene 1-Nonene Nonane 1,9-Decadiene 1-Decene Decane Indene 1,10-Undecadiene 1-Undecene Undecane 1,11-Dodecadiene 1-Dodecene Dodecane 1,12-Tridecadiene 1-Tridecene Tridecane 1,13-Tetradecadiene 1-Tetradecene Tetradecane 1-Pentadecene Pentadecane 1-Hexadecene Hexadecane

C5H10 C5H8 C5H6 C6H12 C6H12 C6H14 C6H8 C6H6 C7H14 C7H16 C7H8 C8H16 C8H16 C8H18 C8H14 C9H16 C8H10 C8H8 C9H18 C9H20 C10H18 C10H20 C10H22 C9H8 C11H20 C11H22 C11H24 C12H22 C12H24 C12H26 C13H24 C13H26 C13H28 C14H26 C14H28 C14H30 C15H30 C15H32 C16H32 C16H34

70 68 66 84 84 84 80 78 98 100 92 112 112 114 110 124 106 104 126 128 138 140 142 116 152 154 156 166 168 170 180 182 184 194 196 198 210 212 224 226

Alkene 1-Alkene Alkadiene Alkadiene iso-Alkane 1-Alkene Alkane Alkadiene Aromatic 1-Alkene Alkane Aromatic iso-Alkane 1-Alkene Alkane iso-Alkane iso-Alkane Aromatic Aromatic 1-Alkene Alkane Alkadiene 1-Alkene Alkane Aromatic Alkadiene 1-Alkene Alkane Alkadiene 1-Alkene Alkane Alkadiene 1-Alkene Alkane Alkadiene 1-Alkene Alkane 1-Alkene Alkane 1-Alkene Alkane

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Energy & Fuels

Table 4. Percentage hydrocarbon distribution for pyrolysis products. Hydrocarbons Alkanes (%) n-Alkanes iso-Alkanes Cycloalkanes Alkenes (%) 1-Alkenes iso-Alkenes Cycloalkenes Propene Butene Alkadienes (%) Aromatics (%)

600 oC

700 oC

800 oC

35.25 27.82 6.68 0.75 61.59 39.70 5.69 1.56 6.62 8.02 2.02 1.13

19.00 15.04 3.64 0.32 65.49 39.05 8.89 4.78 5.54 7.23 10.89 4.62

14.57 11.34 2.92 0.31 59.55 33.05 8.81 6.42 4.62 6.65 15.87 10.00

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Table 5. Product distribution of VR cracking over different catalysts. Catalysts

Y-based catalyst

Silica sand

Calcium aluminate

Calcium aluminate

Temperature (oC)

600

650

600

650

Catalyst-to-oil (g/g) Steam-to-oil (g/g)

8 0.5

8 0.5 Product distribution (wt %) 29.81 5.46 3.09 7.55 1.25 6.96 0.61 4.38 63.36 61.95 19.35 17.40 16.76 8.44 8.24 91.56

8 0.5

8 0.5

22.82 4.24 2.04 7.72 0.45 5.43 0.15 2.05 66.60 70.80 26.47 22.17 18.01 4.16 6.38 95.84

43.69 7.02 3.81 14.49 0.87 11.68 0.29 4.63 70.49 49.88 21.88 14.69 11.02 2.30 6.43 97.70

Cracking gas mathane ethane ethylene propane propylene butanes butenes Olefinicity a Liquid yield b Gasoline Diesel VGO Heavy oil Coke Conversion a

39.84 6.74 3.00 3.45 2.72 9.59 5.10 8.64 54.44 46.87 29.75 13.28 3.66 0.19 13.29 99.81

y(C2=+ C3=+ C4=)/y(cracking gas); b The total yield of gasoline, diesel, VGO and heavy oil.

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Energy & Fuels

Table 6. Gas composition of coke gasification over different catalysts. Catalysts Gasifying agent Temperature (oC) Composition (vol %) H2 CO CH4 CO2 C2 hydrocarbons Carbon conversion (%)

Silica sand

Calcium aluminate

Calcium aluminate

Calcium aluminate

steam 800

steam 780

steam 800

steam+2%O2 800

54.19 9.85 3.22 31.64 1.10 50.36

57.28 5.67 0.27 35.20 1.58 82.46

58.95 5.95 0.35 33.84 0.91 89.79

39.36 5.51 1.57 52.05 1.51 98.69

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Figures Figure 1. Schematic diagram of the coupling process of heavy oil millisecond pyrolysis and coke gasification. Figure 2. Schematic diagram of the fluidized bed experimental apparatus. Figure 3. Effect of the residence time on the coke yield. Figure 4. Effect of reaction temperature on the residue yield. Figure 5. Effect of reaction temperature on the reaction rate. Figure 6. Py-GC/MS chromatograms from millisecond pyrolysis of Shengli VR. Figure 7. The main reaction pathways for the millisecond pyrolysis of VR. Figure 8. Effects of pyrolysis temperature on the peak area of alkenes, alkanes, alkadienes and aromatics. Figure 9. Effects of pyrolysis temperature on the peak area ratio of alkenes to alkanes. Figure 10. Effects of catalyst on the gasification of coke (TG analysis with the heating rate of 10 °C/min under steam atmosphere). Figure 11. XRD spectra characterization of the fresh, aged (with steam at 800°C for 8h) and regenerated (reused for 3 times) calcium aluminate catalysts.

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Energy & Fuels

Figure 1. Schematic diagram of the coupling process of heavy oil millisecond pyrolysis and coke gasification.

39


Energy & Fuels

Gas sampling

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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Figure 2. Schematic diagram of the fluidized bed experimental apparatus.

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Energy & Fuels

Figure 3. Effect of the residence time on the coke yield.

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Figure 4. Effect of reaction temperature on the residue yield.

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3.5 r700 C/r600 C o

o

r800 C/r700 C

3.0

o

o

2.5 r/r

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

2.0 1.5 1.0 0

5

10

15

20

Reaction time (s) Figure 5. Effect of reaction temperature on the reaction rate.

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Energy & Fuels

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Figure 6. Py-GC/MS chromatograms from millisecond pyrolysis of Shengli VR.

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Energy & Fuels

Aromatics

cracking

Coke

condensation

aromatization aromatization

VR cracking

Alkenes cracking Alkadienes cracking

cracking

Alkanes

Figure 7. The main reaction pathways for the millisecond pyrolysis of VR.

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Energy & Fuels

6

7

7x10

o

(a)

600 C o 700 C o 800 C

o

(b)

600 C o 700 C o 800 C

6

6x10 Peak area of alkanes

Peak area of alkenes

3x10

7

2x10

7

1x10

6

5x10

6

4x10

6

3x10

6

2x10

6

1x10 0

0 3

4

5

6

7

8

9 10 11 12 13 14 15 16

4

5

6

Carbon number

7

8

9 10 11 12 13 14 15 16 Carbon number

6

8x10

o

600 C o 700 C o 800 C

(c)

7

1.5x10

o

(d)

6

600 C o 700 C o 800 C

7x10 Peak area of aromatics

7

2.0x10 Peak area of alkadienes

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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7

1.0x10

6

5.0x10

6

6x10

6

5x10

6

4x10

6

3x10

6

2x10

6

1x10 0.0

0 4

5

6

7

8

9 10 11 12 13 14 15 16

6

7

Carbon number

8

9

10

11

12

Carbon number

Figure 8. Effects of pyrolysis temperature on the peak area of alkenes, alkanes, alkadienes and aromatics ((a) peak area of alkenes; (b) peak area of alkanes; (c) peak area of alkadienes; (d) peak area of aromatics).

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50 45 Peak area ratio (Alkene/Alkane)

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Energy & Fuels

o

600 C o 700 C o 800 C

25 20 15 10 5 0 4

5

6

7

8

9 10 11 12 13 14 15 16 Carbon number

Figure 9. Effects of pyrolysis temperature on the peak area ratio of alkenes to alkanes.


Energy & Fuels

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Figure 10. Effects of catalyst on the gasification of coke (TG analysis with the heating rate of 10 °C/min under steam atmosphere).


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5000

Regenerated

4000 3000 2000 1000 50000

Aged

4000

Intensity

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

3000 2000 1000 50000

Fresh

4000 3000 2000 1000 0

10

20

30

40

50

60

70

Angle Figure 11. XRD spectra characterization of the fresh, aged (with steam at 800°C for 8h) and regenerated (reused for 3 times) calcium aluminate catalysts.


Coupling Process of Heavy Oil Millisecond Pyrolysis and Coke

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