Efficient Conversion of Light Cycle Oil into High ... - ACS Publications

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Efficient Conversion of Light Cycle Oil into High-Octane-Number Gasoline and Light Olefins over a Mesoporous ZSM‑5 Catalyst Li Xin, Xinxin Liu, Xiaobo Chen, Xiang Feng, Yibin Liu, and Chaohe Yang* State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Qingdao, Shandong 266580, People’s Republic of China S Supporting Information *

ABSTRACT: Producing high-octane-number (ON) gasoline and light olefins is a promising route to valorize light cycle oil (LCO). In this work, the LCO was mildly hydrogenated and then catalytically cracked to produce high-ON gasoline and light olefins. Mesoporous ZSM-5 zeolite (meso-ZSM-5) was prepared and, for the first time, was applied in this process to crack the hydrogenated LCO (hydro-LCO). The catalytic performance of meso-ZSM-5 was evaluated in detail under different reaction temperatures and weight hourly space velocities (WHSVs). The results showed that, in comparison to less than 64 wt % hydroLCO conversion over the conventional ZSM-5 catalyst, the novel catalyst exhibited excellent performance in cracking hydroLCO with quite a high conversion of 84.8 wt %, affording a gasoline yield of 56.4 wt % and light olefin yield of 19.3 wt % at 560 °C and 10 h−1. In addition, the conversion behaviors of hydro-LCO components were analyzed over both the conventional ZSM-5 and meso-ZSM-5 catalysts. Finally, on the basis of the study of the acid and pore properties of both catalysts, a detailed intrinsic reason for enhanced performance was elucidated. It demonstrated that the remarkable catalytic performance of the meso-ZSM-5 catalyst was closely related to the high diffusion of reactants and the accessibility of acid sites.

1. INTRODUCTION Fluid catalytic cracking (FCC) is a key unit in a refinery to convert heavy oil into light products, such as liquefied petroleum gas (LPG), gasoline, and light cycle oil (LCO). The LCO has been used as a blending component for diesel pool. However, one of the characteristics of LCO is the high content of aromatics, especially polycyclic aromatic hydrocarbons (PAHs). It has been well-known that PAHs in diesel would decrease the cetane number (CN)1−3 and lead to a high level of particulate emissions in exhausts.4−6 With the drive of environmental regulations, the blend of LCO into diesel is severely restricted.7,8 Normally, this stream is disposed as a viscosity cutter for heavy residual fuel with poor value.7,9 Several routes have been proposed to valorize LCO, such as hydrotreating,1,10,11 hydrocracking,12−14 and selective ring opening (SRO).7,15−17 However, the CN improvement by aromatic saturation in the hydrotreating process is very limited.15,18 Hydrocracking can produce alkyl-aromatics from PAHs in LCO. Nevertheless, aromatic oversaturation was usually accompanied, which leads to the low utilization efficiency of hydrogen and considerable yield of light paraffins with a low value.7,19 Moreover, this route involves very high investment and operating costs during industrial application.20 SRO route has been shown to convert LCO into high CN diesel components; however, the hydrogen consumption of this route is very high, and the activity and stability of catalysts for this route still need further improvement.21,22 In short, these routes have drawbacks in the valorization of the LCO stream. It is worth noting that PAHs are easily partially saturated to naphtheno-aromatics by mild hydrogenation, which can be further catalytically cracked over acid sites to generate alkylaromatics and light olefins/alkanes by the opening of the naphthenic ring.19,23,24 In this route, a moderate amount of © XXXX American Chemical Society

hydrogen is consumed and mild operation conditions are involved, while valuable high-octane-number (ON) gasoline and LPG components can be obtained. A similar synergistic process has been proposed in patents by some companies, such as the ExxonMobil Research and Engineering Company,25,26 SK Innovation Co., Ltd.,27,28 and JX Holding, Inc.29 Recently, Jin et al.19 tested this promising route by catalytic cracking of hydrogenated LCO (hydro-LCO). They found that hydroLCO showed good potential to produce high-ON gasoline. However, the conversion of hydro-LCO is below 65 wt % in their study. Zhang et al.30 also performed the catalytic cracking of hydro-LCO under the catalysis of both ZSM-5 and Y zeolite. However, the conversions of hydro-LCO were below 45 and 60 wt % over ZSM-5 and Y catalysts, respectively. This is mainly because the naphtheno-aromatics with a high content in hydro-LCO easily suffer from the hydrogen transfer reaction to form naphthalenes in the Y catalyst, and the severe restriction exists as these two-ring molecules diffuse in the pores of ZSM-5 zeolite.23 It seems that, in the conventional catalytic approach, high catalytic conversion of hydro-LCO is difficult to achieve. In recent years, the mesoporous ZSM-5 (meso-ZSM-5) catalyst has been used for the catalytic cracking process.31−37 For instance, Zhao et al.36 applied meso-ZSM-5 in the catalytic cracking heavy oil and found the improved light olefin yield. Graca et al.37 performed the co-cracking of bio-oils and traditional FCC feedstocks over the meso-ZSM-5 catalyst, and enhanced catalyst deactivation resistance was observed. However, the research on the application of the meso-ZSM-5 catalyst in the LCO conversion is lacking. Received: March 28, 2017 Revised: June 1, 2017

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DOI: 10.1021/acs.energyfuels.7b00852 Energy Fuels XXXX, XXX, XXX−XXX

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2.4.2. Hydro-LCO Catalytic Cracking. The catalytic cracking of the hydro-LCO was performed in a microactivity test (MAT) unit. Before the test, 6 g of catalyst was purged with a flow of N2 for 0.5 h to remove absorbed water at the reaction temperature. Then, 1 ± 0.001 g of hydro-LCO was injected into the reactor by a micropump. After that, the reactor was stripped for 10 min by a nitrogen flow with 40 mL/min. During reaction and stripping steps, the effluent was cooled in an ice−water bath to obtain liquid products. Finally, the gaseous products were collected in a gas bag by the dewatering method. In the catalytic cracking process, the yield of each product was defined as follows:

In this work, we proposed a novel catalyst to efficiently convert LCO into high-ON gasoline and light olefins. Meso-ZSM-5 zeolite was prepared by desilicication and, for the first time, was applied in the synergistic process to convert hydro-LCO. The FCC performance of hydro-LCO over this novel catalyst was evaluated in detail under different reaction temperatures and weight hourly space velocities (WHSVs). Then, the conversion behaviors of hydro-LCO components were analyzed. Furthermore, the acid and pore properties of ZSM-5 and meso-ZSM-5 were studied. On the basis of these, the intrinsic reason for the remarkable catalytic performance of the meso-ZSM-5 catalyst was elucidated. This novel catalyst proposed in this paper shows good potential to efficiently produce high-ON gasoline and light olefins from LCO.

Yi (wt %) =

mass of i in products × 100 mass of hydro‐LCO feed

The hydro-LCO conversion is defined as the sum of the yield of dry gas, LPG (C3 and C4 hydrocarbons), gasoline, heavy oil, and coke.

2. EXPERIMENTAL SECTION

hydro‐LCO conversion (wt %)

2.1. Feedstock. A LCO distillate was used in this study, which was supplied by Jinan Petrochemical Co., Ltd., Sinopec Group. Its properties are listed in Table S1 of the Supporting Information. 2.2. Catalyst Preparation. Meso-ZSM-5 zeolite was prepared by desilicication. NaZSM-5 zeolite (SiO2/Al2O3 = 38), as the starting material, was treated in a NaOH solution with 0.2 mol/L at 80 °C for 6 h. Then, the suspension was filtered and washed. After that, the alkali-treated zeolite went through ion exchange (with NH4NO3 solution of 0.3 mol/L), filtration, and desiccation at 120 °C for 5 h. Finally, the meso-ZSM-5 zeolite was obtained by calcinating these dried filter cakes at 500 °C for 4 h. The conventional ZSM-5 zeolite was obtained by the same process, except for alkali treatment. The catalytic cracking catalyst was prepared as follows: In a typical process, 30 g of zeolite, 65 g of kaolin matrix, and 60 g of water were wellmixed. Then, 5 g of silica sol binder was added. After intensive stirring, the mixture was sprayed to obtain the dried microspheroidal catalyst. 2.3. Catalyst Characterization. Nitrogen adsorption and desorption experiments were performed at −196 °C using a Quadrasorb SI apparatus. The total specific surface area was calculated by the Brunauer−Emmett−Teller (BET) method. The micro- and mesopore surface areas and micropore volume were calculated by the t-plot method. The pore size distribution was obtained by the Barrett− Joyner−Halenda (BJH) method on the adsorption branch of the isotherm. The transmission electron microscopy (TEM) images were acquired by a JEM-2100 microscope at 200 kV. The samples were also characterized by X-ray diffraction (XRD). It was conducted using an X’Pert PRO MPD diffractometer with Cu Kα radiation (λ = 0.154 nm, 40 kV, and 40 mA) with a scanning range of 2θ from 5° to 50° at a scanning rate of 10°/min. The amount of different types of acid sites was determined by Fourier transform infrared spectroscopy (FTIR). Pyridine adsorption was performed on a Bruker Tensor 27 spectrometer by the zeolite adsorption at room temperature for 40 min and further desorption at 200 or 400 °C for 1 h. Afterward, the spectra were recorded at room temperature by accumulating 32 scans at a resolution of 4 cm−1. The bands at 1540 and 1450 cm−1 were used to calculate the amount of Brønsted and Lewis acidity, respectively. The diffusion measurement was performed in an intelligent gravimetric analyzer (IGA, Hiden Analytical, U.K.) using tetralin as a model compound. 2.4. Experimental Process. 2.4.1. LCO Hydrogenation. The diagram of the LCO conversion process is presented in Figure S1 of the Supporting Information. The LCO hydrogenation was carried out in a continuous tubular fixed-bed stainless-steel reactor (8 mm inner diameter and 1000 mm length), heated by a triple-fired furnace. A 5 mL commercial presulfided Ni−Mo/Al2O3 catalyst (its properties are presented in Table S2 of the Supporting Information) was loaded in the middle part of the reactor. During reaction, the feed was mixed with hydrogen and pumped into the reactor. The hydrogenated products were cooled by an ethanol solution at −5 °C in a condenser before entering the gas−liquid separator, in which the hydrogenated oil was collected. The gases were passed through a wet test meter and vented to a tail gas treatment system.

= Ydry gas + YLPG + Ygasoline + Yheavy oil + Ycoke 2.5. Product Analysis. The gaseous products were analyzed on Bruker 456 gas chromatography (GC) equipped with two thermal conductivity detectors (TCDs, for analyzing H2 and N2) and a flame ionization detector (FID, for analyzing C1−C6 hydrocarbons). The hydrocarbon structure analysis of LCO, hydro-LCO, and liquid products of the catalytic cracking process were performed by comprehensive two-dimensional gas chromatography coupled to time-offlight mass spectrometry (GC × GC−TOF MS), and the corresponding content analysis was performed by GC × GC−FID (LECO Corporation, St. Joseph, MI, U.S.A.). The GC × GC system was equipped with an apolar HP-PONA column and a weak polar DB-17HT column. The liquid product of the cracking process was also analyzed by simulated distillation (ASTM D2887) on Varian CP-3800 GC to determine the yields of gasoline (from initial boiling point to 204 °C), unconverted LCO (from 205 to 350 °C), and heavy oil (above 350 °C). The coke content deposited in the catalyst was determined by a high-frequency infrared carbon analyzer.

3. RESULTS AND DISCUSSION 3.1. Mild Hydrogenation of LCO. The mild hydrogenation of LCO was carried out at the following typical conditions: 360 °C, 8.0 MPa, H2/oil (v/v) = 800, and liquid hourly space velocity (LHSV) = 0.7 h−1. Table 1 shows the Table 1. Mass Balance for the Hydrogenation Process product

yield (wt %)

gases liquid coke loss

0.25 99.7 0.001 0.05

mass balance of the hydrogenation process. A remarkably high liquid product yield (99.7 wt %) was obtained in this process. The GC analysis showed that the main gas compositions, apart from H2, were CH4, C2H6, and C3H8, which should be formed by cleavage of C−C bonds over the acid sites of the Ni−Mo/ Al2O3 catalyst. Table 2 presents the hydrocarbon distribution of LCO before and after hydrogenation. It can be seen that the PAH content in the LCO reached up to 60 wt %. As a result, this LCO cut involved a very low cetane index (CI) and a higher density (Table S1 of the Supporting Information). After hydrogenation, the PAHs in LCO were almost converted, of which the content sharply decreased from 60.7 to 3.2 wt %. In comparison of the B


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Energy & Fuels Table 2. Hydrocarbon Distribution of LCO and Hydro-LCO content

a

hydrocarbon distribution (wt %)

before

after

alkanes cycloalkanes olefins alkyl-benzenes (alkyl-bz) tetralins di-aromatics tetrahydro-phenanthrens (tetrahydro-ph) tri-aromatics PAHs (sum)a

28.5 5.2 2.3 2.6 0.7 35.6 10.0 15.1 60.7

30.3 14.8 0 6.9 44.8 1.7 1.2 0.4 3.2

PAHs (sum) = the sum of tetrahydro-ph and di- and tri-aromatics. Figure 2. Product distribution of hydro-LCO catalytic cracking as a function of the reaction temperature.

hydrocarbon distribution of LCO and hydro-LCO, most of the PAHs were selectively saturated to form tetralins. The increment of tetralins was significantly much higher than that of cycloalkanes, indicating that the partial hydrogenation of PAHs to mononuclear aromatics is easy, while the complete saturation to cycloalkanes is difficult.38−40 3.2. FCC Performance of Hydro-LCO over Meso-ZSM-5 and ZSM-5 Catalysts. 3.2.1. Hydro-LCO Conversion under Different Reaction Temperatures over Meso-ZSM-5 and ZSM-5 Catalysts. Figure 1 presents hydro-LCO conversion and

meso-ZSM-5 and ZSM-5 catalysts. It can be seen that the gasoline was the highest yield product on both catalysts. The yield of gasoline was as high as ∼60−52 wt % on the mesoZSM-5 catalyst, which was much higher than that on the ZSM-5 catalyst, with ∼37−31 wt % yield. In addition, the meso-ZSM-5 catalyst showed a LPG yield of 20.4−27.4 wt %, which was also higher than that obtained on the ZSM-5 catalyst with a yield of 9.2−20.9 wt %. Besides, the heavy oil were formed in a very small amount on both catalysts, with less than 0.4 wt % yield. The meso-ZSM-5 catalyst exhibited much higher hydro-LCO conversion; however, its dry gas and coke yields were comparable to those of the ZSM-5 catalyst. This may be due to the mesopores in the meso-ZSM-5 zeolite being in favor of the rapid diffusion of products and reducing their secondary reactions to generate coke and dry gas.41 With the temperature increasing from 520 to 580 °C, both yields of gasoline on meso-ZSM-5 and ZSM-5 catalysts decreased, which should be due to the overcracking of gasoline. As a result, the yield of LPG increased. The dry gas yield also showed increasing trends on both catalysts. This should be for the reason that a high temperature promotes the thermal cracking reactions, characterized by the formation of H2 and C1 and C2 hydrocarbons.42,43 Figure 3 shows light olefin and light alkane yields at different reaction temperatures. The light olefin is dominated in LPG in both catalysts, and the ratio of light alkanes to light olefins is below 0.40, as presented in Table S3 of the Supporting

Figure 1. Hydro-LCO conversion and total yield of LPG and gasoline as a function of the reaction temperature.

total yield of LPG and gasoline as a function of the reaction temperature at a WHSV of 10 h−1. The conversion of hydro-LCO on the meso-ZSM-5 catalyst was as high as 83.3−85.4 wt %. In comparison, the hydro-LCO conversion on the ZSM-5 catalyst was only 49.2−57.7 wt %. This indicates that the meso-ZSM-5 catalyst has remarkable catalytic activity for the hydro-LCO conversion compared to the ZSM-5 catalyst. Figure 1 also shows the total yield of LPG and gasoline on both catalysts. Notably, this total yield on the meso-ZSM-5 catalyst was extremely high, with a range of 79.2−80.3 wt %. However, the total yield of LPG and gasoline on the ZSM-5 catalyst was in a much lower range of 46.4−51.5 wt %. In Figure 1, the gaps between hydro-LCO conversions and total yields of LPG and gasoline were narrow on both catalysts. This demonstrates the high selectivity of LPG and gasoline when the hydro-LCO was catalytically cracked in the MFI structure zeolites. Figure 2 shows the effect of the reaction temperature on the product distribution of hydro-LCO conversion on the

Figure 3. Light olefin and light alkane yields in LPG as a function of the reaction temperature. C


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Energy & Fuels Information, indicating the weak hydrogen transfer reactions as the hydro-LCO cracked in the MFI zeolites. The yield of light olefins on the meso-ZSM-5 catalyst was 14.7−21.4 wt %, which was 7−8 wt % higher than that on the ZSM-5 catalyst, with 6.8−14.9 wt % yield. Specially, with the temperature increasing, the yield of light olefins increased significantly in both catalysts. However, the yield of light alkanes on both catalysts increased mildly, especially on the meso-ZSM-5 catalyst, which shows a negligible increase from 5.5 to 6.1 wt %. This may be because a high temperature facilitates the diffusion of light olefins from pores of the zeolite crystallites and reduces the chance of the light olefins converting to light alkanes by hydrogen transfer reactions. These results also suggest that the gasoline overcracking enhanced by a high temperature mainly lead to the formation of light olefins as hydro-LCO was cracked over ZSM-5 structures. 3.2.2. Hydro-LCO Conversion under Different WHSVs over Meso-ZSM-5 and ZSM-5 Catalysts. The catalytic cracking performance of meso-ZSM-5 and ZSM-5 catalysts was also investigated at different WHSVs by varying the feed rate at 560 °C. A low WHSV implies a longer retention time for the feed stream over the catalyst. As seen in Figure 4, on the meso-ZSM-5

Figure 5. Product distribution of hydro-LCO conversion as a function of the WHSV.

gasoline and LPG were still the main products of the hydroLCO cracking, especially on the meso-ZSM-5 catalyst. In detail, the gasoline yield on the meso-ZSM-5 catalyst peaked at 15 h−1, with a value of 56.2 wt %, which was much higher than the gasoline yield obtained on the ZSM-5 catalyst (less than 36 wt %). In addition, the yield of LPG on the meso-ZSM-5 catalyst was in the range of 19−25 wt %, which also higher than that on the ZSM-5 catalyst (between 12 and 20 wt %). In terms of other products, quite a low heavy oil (below 0.5 wt %) and comparable yields of dry gas and coke yield were also obtained, which were in the ranges of 3.6−1.1 and 2.8−1.2 wt %, respectively. Figure 6 shows the light olefin and light alkane yields in LPG with the WHSV. Light olefins were also dominant in LPG for

Figure 4. Hydro-LCO conversion and total yield of LPG and gasoline as a function of the WHSV.

catalyst, the hydro-LCO also showed a much higher conversion than on the ZSM-5 catalyst. Specifically, with WHSV increasing, the hydro-LCO conversion on the meso-ZSM-5 catalyst increased before reaching a peak with 84.8 wt % at 10 h−1 and then decreased to 76.4 wt % at 20 h−1. The increase of hydro-LCO conversion may be for the reason that an appropriate increase of WHSV would decrease the hydrogen transfer reaction of tetralins and the formation of naphthalenes. However, with a further increase of WHSV, the catalytic cracking reaction will be sharply suppressed by the significant shortness of the retention time, leading to the decrease of hydro-LCO conversion. For the ZSM-5 catalyst, hydro-LCO conversion monotonically decreased from 63.4 to 43.2 wt % with WHSV. This may be because the weakness of the hydroLCO catalytic cracking reaction as a result of the retention time reduction on this catalyst was quite severe, even though the hydrogen transfer reaction of tetralins can be reduced in the meantime. With the WHSV increasing, the total yields of LPG and gasoline follow the trend of conversions on both catalysts. Similar to Figure 1, the gaps between conversions and these total yields were quite small, indicating the high selectivity of the valuable products at the tested WHSV. Figure 5 presents the product distribution of hydro-LCO conversion at the different WHSVs. Under the varied WHSVs,

Figure 6. Light olefin and light alkane yields in LPG as a function of the WHSV.

both catalysts in the tested WHSV. In particular, the yields of light alkanes and light olefins obtained on the ZSM-5 catalyst were decreased with the WHSV increasing. However, the yields of light olefins on the meso-ZSM-5 catalyst showed a peak yield of 19.2 wt % at 10 h−1. This may be for the reason that the low WHSV (meaning long retention time) increases the possibility of light olefins suffering from unwanted hydrogen transfer reactions on the meso-ZSM-5 zeolite channels, which leads to the high yield of light alkanes at WHSVs of 4 and 7 h−1, as also shown in Figure 6. As the WHSV increased to 10 h−1, this possibility reduced and the yield of light olefins increased. However, with a further increase of WHSV from 10 to 20 h−1, D


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reactions of tetralins to di-aromatics, which follow the radical mechanism, will be enhanced.47−49 This may contribute to the increasing yields of di-aromatics as shown in Figure 7a. In addition, conversions of alkyl-benzenes were limited on both catalysts in view of their original content of 6.9 wt % in the hydro-LCO (Table 2), and they were the considerable component in the unconverted hydro-LCO on the meso-ZSM-5 catalyst. This might be for the reason that these alkyl-benzenes mainly involved compounds with single-ring aromatics attaching to multi-short chains, which were relatively unreactive during catalytic cracking. It was worth noticing that the biggest difference in the composition of unconverted hydro-LCO is in tetralins. On the meso-ZSM-5 catalyst, the yield of tetralins was extremely low in the unconverted hydro-LCO. However, this yield was high (even over 22 wt %) on the ZSM-5 catalyst. As a result, tetralins were the dominated component in the unconverted hydro-LCO over the ZSM-5 catalyst. In fact, apart from suffering from the ring-opening reaction to form alkyl-aromatics, the tetralins, as the typical hydrogen donors,50−52 are prone to donate hydride ions to acceptors and generate PAHs. Scheme S1 of the Supporting Information shows the reaction pathways of conversion of tetralins. The detailed conversion of tetralins and their ring-opening selectivity (the definitions are listed in the Supporting Information) at varied reaction temperatures and WHSVs are presented in Figure 8. It is clear that the meso-ZSM-5 catalyst exhibited amazing catalytic activity in the tetralins, in which nearly 100 wt % of these compounds were converted in both panels of Figure 8. In comparison, on the ZSM-5 catalyst, conversion of tetralins was quite lower, with 37.5−62.4 wt % under the tested conditions. As far as the ring-opening selectivity, the meso-ZSM-5 catalyst exhibited a high value of 79−84 wt % at 520−580 °C and 67−85 wt % at 4−20 h−1, respectively. However, the ringopening selectivity on the ZSM-5 catalyst was much lower, which was 45−63 and 52−64 wt % in these tested temperatures and WHSVs, respectively. 3.4. Intrinsic Reason for the High Performance of the Meso-ZSM-5 Catalyst on the Catalytic Cracking of Hydro-LCO. From the experimental results presented above, the high performance of the meso-ZSM-5 catalyst on the hydro-LCO transformation is mainly attributed to the high activity in conversion of tetralins and excellent ring-opening selectivity. In this section, we elucidate the intrinsic reason for the high catalytic performance of meso-ZSM-5 zeolite in converting the tetralins. To identify the pore properties of ZSM-5 and meso-ZSM-5, nitrogen adsorption−desorption experiments were performed. The detailed results are listed in Table 3. As seen, alkali-treated ZSM-5 zeolite (meso-ZSM-5) possessed a larger surface area than its parent zeolite. It should be noted that the mesopore area largely increased, from 19 to 95 m2/g, after alkali treatment. In the meantime, the micopore area decreased, which would be due to the collapse of some lattice in the alkali-treated process. For the pore volume, a significant improvement in the mesopore volume could be observed, which increased from 0.047 to 0.163 cm3/g. Furthermore, the TEM results of ZSM-5 and meso-ZSM-5 in Figure S3 of the Supporting Information also indicated the mesopore formation. In general, this formation of mesopores is explained by the partial extraction of silicon from the zeolite framework (desilication) during the alkali treatment.53−55 In addition, the micopore volume of ZSM-5 decreased by less than 15% (from 0.181 to 0.157 cm3/g) after

the yields of light olefins dropped, owing to severe suppression on the catalytic cracking reaction by the sharp shortness of the retention time. From the experimental studies shown above, it is clear that the meso-ZSM-5 catalyst exhibits high performance in the catalytic cracking of hydro-LCO. Both the high hydro-LCO conversion and high yields of LPG and gasoline were achieved. Moreover, light olefins were the dominated components in LPG, and the high aromatic content in gasoline presented in Figure S2 of the Supporting Information implied the high ON of as-yielded gasoline over the meso-ZSM-5 catalyst. 3.3. Conversion Behaviors of Hydro-LCO Components over Meso-ZSM-5 and ZSM-5 Catalysts. Figure 7 presents

Figure 7. Hydrocarbon distributions of the unconverted hydro-LCO as a function of the (a) reaction temperature and and (b) WHSV.

the composition of unconverted hydro-LCO under different reaction temperatures and WHSVs. It can be seen that, on the whole, the alkanes and cycloalkanes in hydro-LCO were fairly converted over both catalysts because of their fine crackabilities. More specifically, these yields of unconverted alkanes and cycloalkanes on the ZSM-5 catalyst were higher than those on the meso-ZSM-5 catalyst, suggesting that meso-ZSM-5 has more catalytic activity for these saturated structures. Furthermore, in both panels a and b of Figure 7, the unconverted hydro-LCO was dominated by PAHs on the meso-ZSM-5 catalyst. It demonstrates that PAHs are the refractory hydrocarbons and are stable in the catalytic cracking process, which is consistent with the results obtained by other researchers.44−46 As the temperature increases, the direct dehydrogenation E


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Figure 8. Conversion of tetralins and their ring-opening selectivity in tested conditions.

Table 3. Pore Properties of ZSM-5 before and after Alkali Treatment sample

BET surface (m2/g)

micropore area (m2/g)

mesopore area (m2/g)

micropore volume (cm3/g)

mesopore volume (cm3/g)

ZSM-5 meso-ZSM-5

372 392

353 297

19 95

0.181 0.157

0.047 0.163

different acid types are presented in Table 4, obtained by FTIR experiments. It clearly showed that, after the alkali treatment, the weak and strong Brønsted acid amounts decreased and the weak and strong Lewis acid amounts increased. Moreover, the total weak acid amount was equal, while the total strong acid amount obviously decreased, which was consistent with the results obtained from NH3-TPD profiles. Generally, the decrease of acid sites, especially the Brønsted acid sites, would decrease the cracking ability of the catalyst.43,56 The NH3-TPD and pyridine infrared (IR) results showed that the Brønsted acid amount significantly decreased in meso-ZSM-5 and the total acid amount also decreased. These should result in a reduction in the catalytic cracking activity. However, in our studies, the hydro-LCO conversion over meso-ZSM-5 is much higher than that over ZSM-5. Therefore, we would attribute the high performance of the meso-ZSM-5 catalyst to the improvement of its pore properties. The micropore of the ZSM-5 zeolite is in maximum size of 5.6 × 5.3 Å.57 However, in this study, tetralin is in the size of 6.86 × 5.01 Å.23 When the attached alkyl group is taken into account, the size of generic tetralin molecules is larger than the micropore size of ZSM-5 zeolite. Consequently, great diffusion restriction exists when tetralins are cracked over the ZSM-5 zeolite.23 Tetralins could only be cracked over the limited acid sites in the pore mouth or external surface of the ZSM-5 zeolite, which leads to a low conversion of tetralins. However, after the mesopores are introduced to the ZSM-5 zeolite, the tetralins can penetrate into these pores. The acid sites in the internal surface are accessible to these compounds in the catalytic cracking process. Therefore, much higher conversion of tetralins can be obtained over the meso-ZSM-5 catalyst. Figure S5 of the Supporting Information presents the IGA results of tetralin adsorption on both zeolites. The meso-ZSM-5 zeolite exhibited much higher uptakes than the ZSM-5 catalyst, indicating that tetralin diffused into the mesopores of mesoZSM-5. This confirmed the improved diffusion ability of tetralins on meso-ZSM-5.

alkali treatment, suggesting that the crystal structure was well-preserved. This result can also be obtained by the XRD patterns (Figure S4 of the Supporting Information), which showed that all of the diffraction peaks assigned to the ZSM-5 zeolite were preserved in the meso-ZSM-5 zeolite and only a few of these peaks were decreased in diffraction intensity. In short, abundant mesopores were successfully introduced to the ZSM-5 zeolite by alkali treatment, while the micropores were in good preservation. Figure 9 shows the temperature-programmed desorption of ammonia (NH3-TPD) profiles of ZSM-5 and meso-ZSM-5

Figure 9. NH3-TPD profiles of ZSM-5 and meso-ZSM-5 zeolites.

zeolites. The low-temperature region (from 110 to 320 °C) and high-temperature region (from 320 to 550 °C) represent the weak and strong acids, respectively. It can be seen that mesoZSM-5 has a mild change in the weak acid amount; however, the strong acid amount of meso-ZSM-5 significantly decreases in comparison to that of ZSM-5. The detailed amounts of F


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Energy & Fuels Table 4. Acid Amounts of ZSM-5 and Meso-ZSM-5 Zeolites 200 °C

400 °C

sample

Brønsted acid (μmol/g)

Lewis acid (μmol/g)

sum

Brønsted acid (μmol/g)

Lewis acid (μmol/g)

sum

ZSM-5 meso-ZSM-5

554 461

76 158

630 619

509 322

68 96

577 418

increase of the gasoline yield. However, the increase of the WHSV drops the hydro-LCO conversion and gasoline yield. LPG is dominated by light olefins, especially those obtained over the meso-ZSM-5 catalyst. Overall, a high gasoline yield of 56.4 wt % and light olefin yield of 19.3 wt % at 560 °C and 10 h−1 can be acquired by the meso-ZSM-5 catalyst, while these are less than 37 and 15 wt % obtained over the ZSM-5 catalyst. The high catalytic performance of the meso-ZSM-5 catalyst is attribute to its high activity in conversion of tetralins (∼100 wt %) and excellent ring-opening selectivity (>80 wt %). An intrinsic reason study reveals that the enhanced performance of the meso-ZSM-5 catalyst should be due to its high pore diffusion and acid accessibility. The novel meso-ZSM-5 zeolite catalyst shows great potential to efficiently convert LCO into high-ON gasoline and light olefins. Further improving the catalytic cracking activity together with hydrothermal stability by phosphorus and surface modification will be our future work.

The meso-ZSM-5 zeolite also gave much higher tetralin ring-opening selectivity than the ZSM-5 zeolite. A possible explanation can be proposed as follows: As a result of the diffusion restriction, the tetralins would adsorb in the pore mouth of zeolite before they are catalytically cracked. In the meantime, some saturated hydrocarbons, such as long-chain alkanes, are about to diffuse to the channel of zeolites to be cracked. However, these diffusions would be hindered by the adsorption of tetralins on the pore mouth. Consequently, the saturated hydrocarbons would suffer from the crack in the acid sites of the pore mouth to generate lighter alkanes and olefins. It is well-known that the tetralins are the typical hydrogen donors; therefore, the hydrogen transfer reaction would subsequently occur between the lighter olefins and tetralins. Finally, PAHs and lighter alkanes generate. This process can be illustrated by Scheme 1a. In comparison, over the meso-ZSM-5

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Scheme 1. Possible Conversion Process of Tetralins in (a) ZSM-5 Zeolite and (b) Meso-ZSM-5 Zeolite

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.energyfuels.7b00852. Properties of feedstock (Table S1) and hydrogenation catalyst (Table S2), experimental process diagram (Figure S1), ratio of light alkanes to light olefins in LPG (Table S3), compositions of gasoline of hydro-LCO conversion on the meso-ZSM-5 catalyst (Figure S2), reaction pathways of tetralins in the catalytic cracking process (Scheme S1), definition of the conversion of tetralins and their ring-opening selectivity, TEM results (Figure S3) and XRD patterns (Figure S4) of ZSM-5 and meso-ZSM-5 zeolites, and IGA results of tetralin adsorption (Figure S5) (PDF)

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zeolite, as shown in Scheme 1b, the tetralins can diffuse easily in mesopores of the meso-ZSM-5 zeolite. They are mostly cracked in the internal surface of the zeolite to generate light olefins and alkyl-aromatics. Therefore, higher ring-opening selectivity can be achieved in the meso-ZSM-5 zeolite.

AUTHOR INFORMATION

Corresponding Author

*Telephone/Fax: +86-532-86981718. E-mail: yangch@upc. edu.cn.

4. CONCLUSION The efficient conversion of LCO into high-ON gasoline and light olefins was conducted using a novel catalytic cracking catalyst. Meso-ZSM-5 zeolite was demonstrated for the first time to be an excellent catalyst for catalytic cracking of hydroLCO. In the LCO hydrogenation, the PAHs can be partially hydrogenated to tetralins easily; however, the complete saturation of PAHs is difficult. The hydro-LCO catalytic cracking over mesoZSM-5 shows a much higher conversion of ∼85 wt % compared to that over ZSM-5, with less than 64 wt %. Gasoline and LPG are the main products of the hydro-LCO conversion. The increase of the reaction temperature would lead to the decrease of the gasoline yield and the increase of the LPG yield in both catalysts. Over the meso-ZSM-5 catalyst, a mild increase of the WSHV is in favor of hydro-LCO conversion and results in the

ORCID

Xiaobo Chen: 0000-0001-9180-0190 Chaohe Yang: 0000-0001-7299-5690 Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors acknowledge financial support of the National Natural Science Foundation of China (Grants 21476263 and U1462205) and the Fundamental Research Funds of China for the Central Universities (Grants 14CX06116A and 16CX05010A).

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