A Conceptual Coupled Process for Catalytic Cracking of High Acid

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Kinetics, Catalysis, and Reaction Engineering

A Conceptual Coupled Process for Catalytic Cracking of High Acid Crude Oil Yibin Liu, Zhiyuan An, Hao Yan, Xiaobo Chen, Xiang Feng, Yongshan Tu, and Chaohe Yang Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.9b00520 • Publication Date (Web): 09 Mar 2019 Downloaded from http://pubs.acs.org on March 10, 2019

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A Conceptual Coupled Process for Catalytic Cracking of High Acid Crude Oil Yibin Liu*, Zhiyuan An, Hao Yan, Xiaobo Chen, Xiang Feng, Yongshan Tu, and Chaohe Yang State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Qingdao, Shandong 266580, China

ABSTRACT: To process high acid crude oil by catalytic cracking technology, a coupled process (FCD-FCC) involving fluid catalytic deacidification (FCD) of high acid crude oil over a spent catalyst and fluid catalytic cracking (FCC) of heavy oil over a regenerated catalyst was proposed. Experiments were carried out in a confined fluidized bed reactor to test the feasibility of this process. The results indicate that the spent catalyst can be used for deacidification of high acid crude oil to restrain the overcracking of light fractions. The optimal reaction conditions for the catalytic deacidification of Dar acidic crude oil are a temperature of 460 °C, a catalyst-to-oil ratio of 5 and a spent catalyst with a coke content of 1.26 wt%. Compared with the direct catalytic cracking of Dar crude oil over a regenerated catalyst, the yields of light oil and desired product increased by 6.10 wt% and 6.78 wt% in the FCD-FCC process, respectively, and the olefin content of naphtha decreased by approximately 16 wt%, although the naphtha yield was lowered by 2.9 wt%. This strategy sheds new light on the processing of high acid crude oil.

1. INTRODUCTION Worldwide crude oil is becoming lower quality with the increasing demand and consumption of oil resources. High acid crude oil with a total acid number (TAN) >1.0 mgKOH/g is becoming more crucial in the worldwide crude oil market. In recent years, the


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worldwide high acid crude oil production has increased by 0.3% each year, and the output of high acid crude oil accounts for approximately 20% of the current global crude oil output1. The TAN, which is defined as the consumed amount of KOH in milligrams to neutralize the acidic compounds in one gram of oil, is an indicator of the acidic oxygen-containing compound content in crude oils. Among the acidic oxygen-containing compounds, naphthenic acids are the main acidic substances, while other acidic compounds are present in minor proportions2-8. Hence, “naphthenic acids” refers to all carboxylic acids in crude oil with the general formula RCOOH in petroleum refinery. Naphthenic acids are reactive components related to corrosion in refinery units. Refining facilities (e.g., crude oil distillation columns, furnaces, heat exchangers and transfer pipelines) can be corroded seriously when processing high acid crude oils9-10. To reduce the corrosivity of high acid crude oils, blending high TAN crude oils with low acid ones, neutralizing acids with caustic chemicals, injecting corrosion inhibitors and upgrading materials are common industrial methods to mitigate equipment corrosion. However, these approaches have many limitations. Therefore, removing naphthenic acids is considered an effective method. A number of approaches have been developed to remove naphthenic acids from high acid crude oil. In general, these approaches can be classified in two ways: nondestructive deacidification and destructive deacidification. The first approach, referred to as a physical method for naphthenic acid removal, includes adsorption11, solvent and ionic liquid extractions12-17, and neutralization18. The second approach, known as a chemical method for naphthenic acid removal, involves esterification19-21, thermal decomposition22, and catalytic decomposition2329

. Naphthenic acid is easily decomposed into hydrocarbons, carbon dioxide, carbon monoxide

and water in the presence of a catalyst, especially the zeolite catalyst. Under the same experimental conditions, the catalytic conversion of naphthenic acid over the zeolite catalyst is higher than the thermal conversion and catalytic conversion over basic oxides28.



FCC is a key operation for converting heavy oil to light products in refineries, and vacuum gas oil (VGO) is the conventional feedstock of FCC units. However, many researchers have attempted to process crude oil by a catalytic cracking method to obtain light products (e.g., light olefins, naphtha)30-35. In addition, FCC has been successfully employed to process high acid crude oils in China36-37. Tian et al37 proposed a new FCC technology to process high acid crude oil. The high acid crude oil is injected into a riser at 200 °C to react over a regenerated catalyst at a high temperature. Naphthenic acid is decomposed completely, which not only prevents corrosion of the follow-up device, but also eliminates the need for the deacidification refining process of naphtha and light cycle oil (LCO). However, there are two disadvantages in this process. First, the light fraction overcracks because of the high temperature and the high activity of the regenerated catalyst at the bottom of the riser. Moreover, the catalyst deactivates rapidly due to the high carbon residue in the high acid crude oil, which retards the cracking of heavy fractions and deteriorates the product distribution. In the last decade, various coupled FCC technologies were developed. Wang et al38 proposed a novel conceptual process to catalytically crack residue and reform the naphtha simultaneously. They further developed other similar technologies to treat vacuum residue, vacuum gas oil, coker gas oil and FCC cycle oil.39-41 The two-stage riser FCC (TSRFCC) technology, another coupled FCC technology that has been applied in some subsidiaries of China National Petroleum Corporation (CNPC) in China, has enormous advantages in flexible operation, adaption for different feedstocks, and restraint on the overcracking of intermediate products. This technique has been introduced to in situ upgrade the FCC naphtha, synergistically process the coker gas oil and yield light olefins42-48. Based on the TSRFCC technology, we propose a novel FCC process for processing high acid crude oil in this paper that can reduce the overcracking of light fractions and increase the conversion of heavy fractions. The operating conditions of catalytic deacidification of Dar acidic crude oil and


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further catalytic cracking of heavy oil were investigated in a confined fluidized bed reactor. On the basis of the experimental results, a synergistic process for catalytic cracking of high acid crude oil was proposed, and its performance was analyzed. 2. EXPERIMENTAL SECTION 2.1. FCD-FCC process. The coupled process (FCD-FCC) of fluid catalytic deacidification (FCD) and fluid catalytic cracking (FCC) for processing high acid crude oil was proposed based on the following discussion. In direct catalytic cracking of high acid crude oil, there are two contradictions. One is the contradiction between the catalytic decomposition of naphthenic acid and overcracking of the light fractions. The other is the contradiction between the deep catalytic cracking of the heavy fraction and the rapid deactivation of the catalyst. To reduce the overcracking of the light fractions in the crude oil, moderate reaction conditions (e.g., low temperature, low-activity catalyst) are needed. However, for the deep catalytic cracking of the heavy fraction, harsh reaction conditions such as high temperature, high-activity catalyst and large catalyst-to-oil ratio are favorable. Fortunately, the catalytic decomposition of naphthenic acid over a zeolite catalyst readily occurs, and the moderate reaction conditions are sufficient. Therefore, we can establish two reaction zones to provide the opposite reaction conditions. The first reaction zone, which is operated under the moderate reaction conditions, is utilized to deacidify the high acid crude oil by the spent catalyst and partially remove the carbon residue. The second reaction zone operated under harsh reaction conditions is employed to crack the heavy fraction that is obtained from the first reaction zone. Figure 1 shows a schematic of the FCD-FCC process. In the FCD-FCC process, there are two riser reactors, two disengagers, two fractionators and one regenerator. The high acid crude oil is deacidified in the FCD riser first, and then the heavy fraction is cracked in the FCC riser. The high acid crude oil, which is preprocessed by desalination, is pumped into the bottom of the FCD riser, and contacted with the spent catalyst



from the stripping section of the FCC disengager. Deacidification occurs at the low temperature to remove the petroleum acids and carbon residue. The reacted catalysts and generated oil and gas are separated in the FCD disengager. The separated catalyst is stripped in the stripping zone of the FCD disengager and then regenerated by burning coke in the regenerator. The generated oil and gas are distilled into cracking gas, naphtha, LCO and heavy oil in the FCD fractionator. The heavy oil is introduced into the bottom of the FCC riser, contacted with the regenerated catalyst from the regenerator and cracked at a high temperature. The spent catalyst and generated oil and gas in the FCC riser are separated in the FCC disengager. The separated catalyst is stripped in the stripping zone of the FCC disengager and then introduced into the bottom of the FCD riser to conduct the deacidification reaction. The generated oil and gas are fractionated into cracking gas, naphtha, LCO, heavy cycle oil (HCO) and slurry in the FCC fractionator. The HCO is blended into heavy oil from the FCD fractionator to crack again in the FCC riser. The cracking gases from the FCD fractionator and FCC fractionator are blended for further processing. Naphtha and LCO are similarly treated. The FCD-FCC process has the following advantages:  It reduces the overcracking of the light fraction contained in high acid crude oil because the high acid crude oil contacts the low-activity spent catalyst and reacts at a low temperature.  Some carbon residue is removed in the FCD process to lower the coke yield and decrease the deactivating rate of the catalyst in the FCC process.  The deacidified heavy oil contacts and reacts with a high-temperature regenerated catalyst to increase the conversion of heavy oil. It should be noted that this process is designed for commercial units. In the laboratory, we employ a confined fluidized bed reactor to study this process.


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Flue gas Dry gas

F R A C T I O N A T O R

FCC disengager

FCD disengager

Regenerator

FCD reactor

FCC reactor

LPG

F R A C T I O N A T O R

Naphtha LCO

Slurry HCO

High acid crude oil Pre-lifting steam

Figure 1. Schematic of the FCD-FCC process Table 1. Properties of Dar Crude Oil property API gravity density (20 °C) (kg/m3) kinetic (80 °C) (mm2/s) TANa (mgKOH/g) CCRb (wt%) wax (wt%) UOP K nickel (µg/g) vanadium (µg/g) elemental analysis (wt%) C H S N distillation fractions (wt%) naphtha (< 200 °C) middle distillate (200-350 °C) vacuum gas oil (350-540 °C) vacuum residue (> 540 °C) a

value 24.85 901.0 44.28 3.68 5.11 23.84 11.7 53.7 3.56 86.36 12.74 0.20 0.37 6.06 16.51 29.98 47.45

Total acid number. b Conradson carbon residue.

2.2. Feedstock and Catalyst. Crude oil from the Darfur region in Sudan (Dar crude oil) was provided by CNPC Fuel Oil Co., Ltd., China. Its properties are listed in Table 1. The catalytic cracking regenerated catalyst, which was named LVR-60R with USY zeolite as the active component and obtained from FCC unit in Changqing Petrochemical Company of CNPC, was used in this study. The physicochemical properties of the regenerated catalyst



are given in Table 2. Table 2. Properties of the Catalysts property bulk density (kg/m3) surface area (m2/g) pore volume (cm3/g) micro-activity index metal content, μg/g Ni V Fe Na particle size distribution (vol%) 0-40 μm 40-110 μm ＞110 μm

value 730 151 0.22 72 2172 873 5136 4907 12.0 68.8 19.2

2.3. Experimental Setup. A confined fluidized bed reactor system (Figure S1), which is operated in a fluidized mode, was used to carry out the experiments. The catalyst was put into the reactor before each run. The reactor was heated by an electric furnace, and the temperature was varied by a temperature controller with a thermocouple in the reactor. The catalyst loading quantity was changed from 160 g to 280 g to vary catalyst-to-oil ratio, and the reaction temperature was adjusted from 440 °C to 500 °C. The feedstock was pumped into the preheating furnace to preheat at 350 °C and then into the reactor along with steam. The feed rate was 40 g/min, and the feeding time was 1 min. The feedstock was cracked into gas, naphtha, LCO, and heavy oil or condensed to coke. During the reaction step, the liquid products were collected in a cold trap located at the exit of the reactor, which was kept at -2 °C by a circulating refrigeration pump. The gaseous products were collected in a gas-collecting bottle by water exchange. The coked catalyst was stripped with steam for 20 minutes to remove entrapped hydrocarbons after the reaction and then taken out by a vacuum pump. The coke content was measured by an infrared absorption carbon-sulfur analyzer. The mass balances for all runs were between 95 and 102 wt% of the injected feed.


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2.4. Product Analysis. The gaseous products were measured and then analyzed by a Variance GC-3800 gas chromatograph to determine the volume percentage of H2, N2, CO, and CO2 and the mass percentage of C1-C6 hydrocarbons. The collected liquid products were weighed and then analyzed by a simulated distillation conducted by an Agilent 6890N gas chromatograph according to the ASTM D 2887 method. Based on the product analysis data, the yields of dry gas (H2, C1, C2, CO, CO2), liquefied petroleum gas (LPG, C3 and C4), naphtha (C5-200 °C), LCO (201-350 °C), and heavy oil (＞350 °C) were determined. The coke yield was determined by measuring the carbon content of the stripped catalyst. In this study, conversion of Dar crude oil is defined as the presence of a (＞350 °C) fraction in the crude oil cracked to gas, light oil and coke as given below: xcrude 

mF,>350 -mP,350 mF,350

100% (1)

Besides, deacidification ratio of high acid crude oil is defined as:



TAN F -TAN L  yL 100% (2) TAN F

where xcrude stands for the conversion of crude oil; mF,>350 and mP,>350 refer to weights of the (>350 °C) fractions in crude oil and product, respectively; φ is the deacidification rate, and yL is the yield of liquid product. TANF and TANL represent the TAN of Dar crude oil and the TAN of liquid product, respectively. Conversion of heavy oil is defined as the total yield of dry gas, LPG, naphtha, LCO and coke in catalytic cracking of heavy oil. Light oil refers to the sum of naphtha and LCO, and the desired product contains LPG, naphtha and LCO. In addition, liquid products from the FCD of Dar crude oil include naphtha, LCO and heavy oil. 2.5. Characterization of the Catalyst. Nitrogen adsorption−desorption isotherms were measured at 77 K on a Micromeritics ASAP 2020 instrument, and the surface area and pore



volume data of the catalysts were calculated from Brunauer−Emmett−Teller (BET) and Barrett−Joyner−Halenda (BJH) equations as well as the t-plot method. The total acid amounts of the catalysts were measured by temperature-programmed desorption of ammonia (NH3TPD). The types and concentrations of acid sites were determined by the adsorption of pyridine, and Fourier Transform Infrared spectra were recorded on a Bruker Tensor 27 spectrometer (PyIR) at diﬀerent desorption temperatures. The catalyst microactivity was evaluated in a ﬁxed bed microreactor according to the ASTM D3907 standard. 3. RESULTS AND DISCUSSION 3.1. Operating Conditions for the Deacidification of Dar Crude Oil. In the FCD process, the purpose is to deacidify the high acid crude oil and remove the carbon residue to reduce the coke yield in the FCC process. Simultaneously, overcracking of light fractions in the crude oil should be minimized. In this section, the effects of operating conditions (e.g., coke content of the spent catalyst, temperature, and catalyst-to-oil (CTO) ratio) were investigated. 3.1.1 Coke Content of Spent Catalysts Two spent catalysts with different coke contents (1.26 wt% and 2.08 wt%) were used to study the effect of the coke content of the catalyst on the catalytic deacidification of the high acid crude oil. The spent catalyst with a coke content of 1.26 wt% (spent catalyst-1) was obtained from catalytic cracking of vacuum gas oil of the high acid crude oil at 500 °C, a CTO ratio of 5 and a WHSV of 15 h-1 in the confined fluidized bed reactor. The spent catalyst with a coke content of 2.08 wt% (spent catalyst-2) was produced by repeating the above procedure using spent catalyst-1 as the catalyst. The catalytic deacidification of Dar crude oil over a spent catalyst was conducted at 500 °C and with a CTO ratio of 5. Table 3 presents the conversion and product yields of the catalytic deacidification of Dar crude oil over two spent catalysts. The results show that the spent catalyst-1 still has good catalytic activity. The TAN of the liquid product is only 0.262 mgKOH/g, and the deacidification rate is more than 94%. However, these values become 1.073


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mgKOH/g and 77.1% for the catalytic deacidification of Dar crude oil over spent catalyst-2. Otherwise, the conversion and the yields of the light oil and desired product from spent catalyst1 are higher than those from spent catalyst-2, while the coke yield is smaller. Hence, the spent catalyst-1 is suitable for deacidification of Dar crude oil. Table 3. Conversion and Product Yields of the Catalytic Deacidification of Dar Crude Oil product conversion (wt%) yield of light oil (wt%) yield of desired product (wt%) TAN of liquid product (mgKOH/g) deacidication rate (%) product yields (wt%) dry gas LPG naphtha LCO heavy oil coke

spent catalyst-1 72.85 54.98 67.81 0.262 94.6

spent catalyst-2 47.77 33.02 43.63 1.073 78.6

1.91 12.82 31.47 23.51 21.03 9.26

1.83 10.62 14.45 18.56 40.44 14.10

To further verify the suitability of spent catalyst-1 for the deacidification of high acid crude oil, we tested the microactivity, structural and acid properties of the regenerated catalyst and spent catalysts. Microactivity was evaluated under the ASTM D3907 standard. Compared with the regenerated catalyst, the microactivity of the spent catalysts decreased, with values of 72 for the regenerated catalyst, 56 for spent catalyst-1 and 45 for spent catalyst-2. Spent catalyst-1 has higher microactivity than other regenerated catalyst (e. g., LTB-2)49, so it still has good catalytic activity. N2 adsorption was employed to analyze the structural characteristics of the catalysts. Figure 2(a) presents the N2 adsorption-desorption isotherms of the regenerated catalyst and two spent catalysts. The isotherms indicate that the quantity of N2 adsorbed by spent catalysts decreased and that as more coke was deposited, less N2 was adsorbed due to the blockage50-52. However, the curves still exhibit high adsorption isotherms at a low relative pressure and hysteresis loops at a high relative pressure due to the presence of micropores and mesopores. The isotherms of the spent catalysts are almost the same as that of the regenerated catalyst,



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which indicates that there are only slight changes in the pore structure of the catalysts after the coke deposition. The structural characteristics of the catalysts are listed in Table 4. The surface area and pore volume, in particular those of micropores, decrease, showing that more coke deposits in micropores and has a more obvious effect on the micropore structure. However, for the deacidification of high acid crude oil, the molecular size of the reactants is much larger than the size of the micropore, and the reactions mainly occur on the external surface. The loss rate of the external surface area is negligible. Table 4. Structural Characteristics of Regenerated and Spent Catalysts parameter

regenerated catalyst

BET surface area (m2/g) micropore surface area (m2/g) external surface area (m2/g) micropore volume (cm3/g) pore volume (cm3/g)

151 83 68 0.043 0.218

spent catalyst-1 residual loss rate value (%) 120 20.53 57 31.33 63 7.35 0.028 34.88 0.207 5.05

spent catalyst-2 residual loss rate value (%) 99 34.44 45 45.78 54 20.59 0.024 44.19 0.17 22.02

The acid properties of the regenerated and spent catalysts were investigated by NH3-TPD and Py-IR, and the curves are displayed in Figure 2(b) and Figure 2(c), respectively. The NH3TPD curves were divided into three regions according to the desorption of ammonia from the weak (100-255 °C), medium (255-365 °C), and strong (365-500 °C) acid sites.43 Figure 2(b) shows that the coke deposition on the catalyst changed the acid strength distribution, whereas cokes covered more strong acid sites than weak acid sites. The total acid amounts of the spent catalyst-1 and spent catalyst-2 decreased by 16.4% and 22.0%, respectively, as shown by integrating the area of the NH3-TPD curves. The values of relative acidity loss are much higher than those in the study of Fals et al.53 In Figure 2(c) of the Py-IR spectra, the peaks of the Lewis acid and Brønsted acid decrease with increasing coke content of the catalysts. Nonetheless, the ratio of Lewis acid to Brønsted acid has slight changes, which means that the coke deposition has no selectivity for the acid sites.


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Figure 2. (a) N2 adsorption-desorption isotherms; (b) NH3-TPD curves; (c) IR spectra in the region characteristic of adsorbed pyridine vibrations

3.1.2 Reaction Temperature The reaction temperature is a main process parameter. It changes the conversion and product distribution in the catalytic cracking process. Figure 3 shows the conversion and product yields of the catalytic deacidifiction of Dar crude oil over spent catalyst-1 as a function of reaction temperature at a CTO ratio of 5. The feed conversion increased from 73.18 wt% to 75.35 wt% with an increase in the reaction temperature from 440 °C to 500 °C. With increasing reaction temperature, the yields of dry gas, LPG and coke increased. It is noteworthy that these increases are trivial. Naphtha and LCO have maximum yields at 460 °C. The desired product yield is also maximized at this temperature. When the temperature was increased to 500 °C, the light oil yield was reduced by 3.3 wt%, and the desired product yield was decreased by 4.5 wt%. 50

100

(a)

coke naphtha

heavy oil LPG

LCO dry gas

30

20

10

(b)

conversion liquid product yield light oil yield desired product yield

90

Conversion / Yield (wt%)

40

Yield (wt%)

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


80

70

60

50

0 440

450

460

470

480

490

500

440

450

Temperature (oC)

460

470

480

490

500

Temperature (oC)

Figure 3. Conversion and product yields of the catalytic deacidification of Dar crude oil over spent catalyst-1 as a function of temperature

Generally, lowering the reaction temperature can reduce the cracking reaction, which is beneficial for obtaining more liquid product. However, under the investigated reaction conditions, the liquid product yield at 440 °C is 69.68 wt%, which is 4.5 wt% less than that at 460 °C. Figure 3(a) indicates that the coke yield at 440 °C is very high. This result is because the temperature is too low to fully vaporize the heavy fractions in the crude oil. The



unvaporized liquid phase that adhered to the surface of the catalyst increased the occurence of the coking reaction. Based on the above discussion, the optimal reaction temperature for the catalytic deacidification of Dar crude oil over a spent catalyst is 460 °C. At this temperature, we obtained the highest liquid product yield, the highest light oil yield and the highest desired product yield, and the acid removal rate was high enough. 3.1.3 Catalyst-to-Oil Ratio A high CTO ratio increases the amount of active sites that make contact with oil, thus improving the contact opportunities. Otherwise, a high CTO ratio reduces the average content of the coke on the catalyst at the same coke yield and favors the stability of the catalyst activity. Moreover, in a confined fluidized bed reactor, the WHSV changes along with the variation in the CTO ratio. Increasing the CTO ratio reduces the WHSV, and prolongs the contact time accordingly. In this study, we set the CTO ratio at 4, 5, 6 and 7, respectively. The WHSV decreases from 15 h-1 to 8.6 h-1 correspondingly. The conversion and product yields versus various CTO ratios at the fixed reaction temperature of 460 °C are shown in Figure 4. 50

100

(a)

coke naphtha

heavy oil LPG

(b)

LCO dry gas

30

20

10

conversion light oil yield desired product yield liquid product yield

90 Conversion / Yield (wt%)

40

Yield (wt%)

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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80

70

60

50

0 4

5

6

7

4

5

6

7

CTO

CTO

Figure 4. Conversion and product yields of the catalytic deacidification of Dar crude oil over spent catalyst-1 as a function of the CTO ratio

Even if the activity of the spent catalyst is greatly reduced, the CTO ratio still has a great influence on the product yields. With an increase in the CTO ratio, the feed conversion and LPG yield increased while the naphtha yield decreased. The yields of LCO, light oil and liquid product first increased and then decreased and had maximum values at a CTO ratio of 5. The


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desired product yield increased from 66.68 wt% at a CTO ratio of 4 to 69.57 wt% at a CTO ratio of 5 and then changed imperceptibly. Light oil overcracking to LPG was observed with an increasing of CTO ratio, and the yield of dry gas showed a negligible change. Increasing the CTO ratio does not noly improve the amount of active sites but also prolongs the reaction time. Otherwise, the higher CTO ratio supplies more active sites for all molecules and increases the contact opportunities between intermediates, such as naphtha and LCO, and catalysts. In general, increasing time is favorable for the thermal cracking that leads to the production of dry gas. However, in this study, the reaction temperature was 460 °C, so the contribution of thermal cracking was low. In addition, the TAN of the liquid product is not more than 0.3 mgKOH/g. Hence, we selected a CTO ratio of 5 as the optimal value for the catalytic deacidification of Dar crude oil. 3.2. Coupled FCD-FCC process. In the FCD-FCC process, the heavy oil from the FCD process would be pumped into the FCC riser to crack again. To conduct the coupled process, we carried out the catalytic deacidification of Dar crude oil at 460 °C and a CTO ratio of 5 many times to obtain the liquid product. The heavy oil in the liquid product was separated by true boiling point distillation and then cracked over regenerated catalyst at 500 °C and a CTO ratio of 5. As a comparison, Dar crude oil was also cracked over the regenerated catalyst at 500 °C and a CTO ratio of 5 to conduct the direct FCC of crude oil. The results of the coupled FCD-FCC process were calculated based on eq. (3) ~ eq. (5). For the yield of heavy oil,

yh  yh,D  yh,C (3) For the other products,

yi  yi,D  yi,C  yh,D (4) For the group compositions of the naphtha from the FCD-FCC process,

xi 

yn,D xi,D  yh,D yn,C xi,C yn,D  yh,D yn,C

(5)



where y stands for the yield, and x represents the group content in naphtha; the subscripts h and n refer to heavy oil and naphtha, respectively; the subscripts C and D represent the catalytic cracking of heavy oil and catalytic deacidification of Dar crude oil, respectively. The reaction conditions and product distributions of these two processes are presented in Table S1. Figure 5 shows a comparison of the product yields in these two processes and the group compositions of naphtha. The Dar crude oil was converted efficiently in both processes, and it was converted more in the FCD-FCC process than in the direct FCC process. We obtained more LPG and LCO but less dry gas, naphtha and coke by the FCD-FCC process than by the direct FCC process. However, the yields of light oil and the desired product increased by 6.10 wt% and 6.78 wt% in the FCD-FCC process, respectively, compared to the direct FCC process,. This result indicates that the FCD-FCC process can restrain the overcracking of light oils. Moreover, the olefin content in naphtha from the FCD-FCC process decreased by 16 wt%. The contributions of the catalytic deacidification of Dar crude oil and the catalytic cracking of heavy oil to product yields in the FCD-FCC process are listed in Figure S2. This result illustrates that the catalytic cracking of heavy oil offers 11.28 % naphtha and 18.19 % LCO but 29.04% dry gas and 24.93 % LPG. 50

50 (a) 39.48

40

(b)

direct FCC FCD-FCC

42.39

direct FCC FCD-FCC

45.22

40 34.17 31.77

29.88

30

30

wt %

Yield (wt%)

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

Page 16 of 23

20.87

20

28.09 27.32

20

15.32 14.64

15.21 12.24

11.87

10

8.16 6.21

7.50

10 6.24

1.60 2.07

0

Dry gas

LPG

naphtha

LCO Heavy oil Coke

0

paraffin

naphthene

aromatic

olefin

Figure 5. Comparison of product yields (a) and naphtha compositions (b) for the direct FCC and the FCD-FCC of Dar crude oil

4. CONCLUSION A coupled FCD-FCC process with catalytic deacidification of high acid crude oil over a spent catalyst and catalytic cracking of heavy oil over a regenerated catalyst was proposed to


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convert high acid crude oil into more LPG, naphtha and LOC, and the experiments were conducted in a confined fluidized bed reactor. Catalytic deacidification of high acid crude oil readily achieves over catalytic cracking catalysts, even spent catalysts. Based on the characterizations and product yields, a spent catalyst with a coke content of 1.26 wt% was used to optimize the reaction conditions for the catalytic deacidification of Dar acidic crude oil. Under the optimal conditions with a temperature of 460 °C and a CTO ratio of 5, we obtained the deacidified liquid product with a yield of more than 80 wt%. A comparison between the direct catalytic cracking of Dar crude oil over the regenerated catalyst and the FCD-FCC process was also carried out in this study. More light oil and desired products were obtained in the FCD-FCC process of Dar crude oil. The naphtha yield decreased by 2.9 wt% compared with the direct FCC process, but the olefin content decreased by 16 wt%. AUTHOR INFORMATION Corresponding Author *Telephone: +86-532-86980917. E-mail: [email protected]. Notes The authors declare no competing financial interest. ACKNOWLEDGEMENTS This work was supported by Natural Science Foundation of China (21776312, 21606254), Key Research and Development Plan of Shandong Province (2018GGX107005, 2017GSF17126), and Fundamental Research Funds for the Central Universities (18CX02130A, 18CX02014A). We also thank Ms. Suhong Huang for her hard work. SUPPORTING INFORMATION - simplified schematic of the experimental setup - contribution of FCD and FCC to product yields - reaction conditions and product distribution of direct FCC and FCD-FCC of Dar crude oil



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For Table of Contents Only


A Conceptual Coupled Process for Catalytic Cracking of High Acid

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