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Coproduction of Syngas during Regeneration of Coked Catalyst for Upgrading Heavy Petroleum Feeds Gang Tian, Gang Wang,* Chunming Xu, and Jinsen Gao State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Beijing 102249, China ABSTRACT: The gasification reactivity of coke, deposited on catalyst during upgrading heavy petroleum feeds in a FCC-like process, with steam and CO2 at 800−900 °C, were investigated. There exist large differences between the performances of catalyst regeneration under steam and CO2 atmospheres. For example, the reactivity of coke-on-catalyst (COC) during steam gasification is about 3−5 times higher than that during CO2 gasification. And the removal rate decreased significantly when conversion was greater than 0.5 during CO2 gasification. In accordance with these results, selecting steam as gasifying agent, a conceptual process of regeneration of coked-catalyst is presented, where the syngas is produced by removing the coke deposited on catalyst by a combination of coke combustion with oxygen and gasification with steam. The performances of COC combustion with oxygen and gasification with steam could be well-described by the homogeneous model and shrinking-core model. Diffusion effects are unobvious during oxygen combustion but apparent during steam gasification when the conversion is high. Additionally, the value of activation energy is about 102 kJ/mol for oxygen combustion and 216 kJ/mol for steam gasification.

1. INTRODUCTION Upgrading heavy petroleum feeds into light products is crucial to satisfy increasing energy demands. Therefore, a number of technical methods, such as residue fluid catalytic cracking (RFCC), hydrocracking, and hydroprocessing, have been developed to process heavy feeds on the basis of carbon rejection and hydrogen addition.1−6 With these processes to upgrading heavy feedstock, one of the challenges is to handle contaminants in feeds properly. These impurities (in particular, metals and Conradson carbon residue (CCR)) can cause many problems in conversion units (such as vanadium and nickel) and act as poisons for catalysts; in addition, CCR causes deactivation of catalysts.7−10 In the previous work, we have reported a study for upgrading the residue.11,12 This scheme uses solid particles with low cracking activity to upgrade heavy feeds in a fluidized-bed reactor. Asphaltenes and most of metals are deposited on the solids; as a result, CCR values and the metal content of the residues are reduced to tolerable levels before further processing. During this process, a substantial amount of coke (carbonaceous compounds) is deposited on the catalyst, and then the catalyst deactivates. Coke combustion is generally employed to recover the spent catalyst activity. Combustion in the regenerator is an exothermic reaction that provides the heat necessary for the cracking process. Previous results show that coke yield exceeds 15 wt % when pretreated feedstock have CCR higher than 20 wt %. Except for the coke needed to maintain unit heat balance, at least 8 wt % coke yields will be burned during catalyst regeneration, which results in the release of excess heat. Generally, part of the heat produced through the combustion of coke is recovered by means of catalyst coolers that control the regenerator temperature and produce steam. It is a low-value way for surplus coke utilization. Besides combustion, coke can be removed from the catalyst surface by gasification with steam or carbon dioxide. Gasification of coke is an endothermic reaction and produces syngas (H2 and © 2013 American Chemical Society

CO). Thus, the heat balance of system can be regulated by coke gasification together with coke combustion in the regenerator. Coproduction of syngas combined with upgrading of heavy petroleum feeds has been receiving attention for decades. Guyer’s13 early work demonstrates that the coke deposited on solid cracking catalyst interacts with a mixture of steam and oxygen can produce a suitable Fischer−Tropsch syngas. Hettinger14 envisions a regeneration system in which CO2 would be utilized in a first stage to remove most of the hydrogen on the coke, and some carbon, and with possibly a second regeneration that would release enough heat to carry out cracking. Moreover, a number of processes mostly based on fluid catalytic cracking (FCC) process, introduce O2/H2O, O2/ CO2, and O2/H2O/CO2 mixtures into regenerator(s) instead of air through one or two steps.15−18 The essence of these processes is, first, partial coke on the catalyst burn with oxygen to release heat, which increases the temperature of particles. Subsequently, most of remaining coke can be gasified with a gasifying agent. In such a way, syngas is produced during regeneration of coked catalyst, meanwhile, transforming lowvalue coke into high-value hydrogen instead of heat. Most of the reported works place more emphasis on patented design; the reactivities of COC gasification with steam or CO2 is seldom mentioned. Comparing with FCC catalysts, the catalysts used in this work are suitable for upgrading residua and then recovered by gasification, because it has a low catalytic activity and excellent resistance to hydrothermal capability and no zeolite in the particles. In addition, most of the work reported in the literature concerning the feeds with CCR is not higher than 10 wt %. However, the Received: Revised: Accepted: Published: 16737

August 2, 2013 October 16, 2013 November 4, 2013 November 4, 2013 dx.doi.org/10.1021/ie402527t | Ind. Eng. Chem. Res. 2013, 52, 16737−16744

Industrial & Engineering Chemistry Research

Article

Figure 1. Schematic diagram of fixed-bed gasification setup.

were conducted in batch mode, with a fixed amount of feedstock and varying time of injection. Coked catalyst was stripped with steam for 30 min to recover entrapped hydrocarbons. More details about this experimental setup schematic and procedure were described elsewhere.11 The density and CCR of LVR was 0.9938 g/cm3 and 17.82 wt %, respectively. The coked catalysts were prepared at 500 °C, and labeled as CLVR, with a carbon content of 2.44 wt %. 2.3. Experimental Apparatus and Procedure. The experiments were conducted in a fixed-bed reactor. A schematic diagram of the experimental setup is shown in Figure 1. The apparatus consisted of an electrical furnace with an internal tube reactor (stainless steel with 0.01 m inside diameter ×0.35 m height), a steam generator, and a gas collection system. The coked catalyst was heated for 30 min at 750 °C under nitrogen flow. Lighter components of coke were consequently removed. These lighter components complicated the acquisition of reliable data. Prior to gasification, 1.0 g of cokedcatalyst “sandwiched” between two quartz wool plugs was placed in the middle of the reactor. Before starting the test, air was removed from the tube reactor using nitrogen. The reactor was then heated to the desired temperature with a ramp at 40 °C/min under nitrogen flow. A proportional-integral-derivative temperature controller was used to control reaction temperature within ±1 °C. When the system reached isothermal steady state and reactant gas was introduced into the reactor, liquid and gas mass flow controllers were used to control the flow rates of the dry air or water and nitrogen, thereby ensuring that the desired steam or oxygen concentration remained constant. The sample was allowed to react until the end of a required reaction time. Afterward, the reactor was cooled to prevent further reactions by injecting nitrogen gas. To eliminate the influence of external diffusion, a series of preliminary tests were performed to determine the proper mass flow rate. The reactivity profiles remained relatively unchanged when water mass flow is higher than 0.4 g/min. Gas film diffusion was negligible at water mass flow rates greater than 0.4 g/min (weight hourly space velocity (WHSV) of 24 h−1). With the same method, the gas film diffusion is eliminated when the O2/N2 is the gas reactant and the flow rates is greater than 150

more inferior heavy oils are used in this study; thus, fundamental properties of coke on the catalyst such as amount, type, and location are different. Therefore, for the sake of application, it is critical that the gasification reactivities of COC under different temperatures are studied, and then a proper method can be schemed. Moreover, the design of the reactor− regenerator system requires the knowledge of the heat released and the kinetics of combustion and gasification. The objective of the present study is to investigate performance of COC under different atmospheres (steam or carbon dioxide) to evaluate COC gasification reactivity. In accordance with the above results, to select a suitable gasifying agent, a conceptual process of coproduction of syngas during regeneration of coked catalyst is proposed, and the related kinetics of this process is also studied.

2. EXPERIMENTAL SECTION 2.1. Catalyst. A fluidizable porous microsphere solid catalyst was used in this study. This catalyst, with affinity properties for contaminants such as asphaltenes and metals, was prepared by the State Key Laboratory of Heavy Oil Processing at China University of Petroleum. It was designed for upgrading residua in a fluidized bed reactor. Moreover, this catalyst has a fairly high pore volume and surface area that contributes to the mesoporous structure but has a low catalytic activity and excellent resistance to hydrothermal capability. The basic physical and chemical properties of the catalyst showed in our previous work.11 In this work, the catalyst particles were subjected to several coking (at 500 °C), steam gasification (at 900 °C) cycles. The results showed that the microactivity kept at about 10, pore volume decreased from 0.277 to 0.267 cm3/g, and pore distribution of the catalysts did not vary significantly after five reaction/regeneration cycles. It was verified that the particle structure can be kept steady during steam gasification. 2.2. Coked Catalyst Preparation. The coked catalysts used in this study were prepared from Chinese Liaohe vacuum residue (LVR) pretreatment experiments conducted in a fluidized bed reactor. The reactor included three zones from top to bottom: the dilute phase zone, the main reaction zone, and the contact zone of feedstock and catalyst. Experiments 16738

dx.doi.org/10.1021/ie402527t | Ind. Eng. Chem. Res. 2013, 52, 16737−16744


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Figure 2. (a) X−t curves for COC during steam gasification at different temperatures. (b) X−t curves for COC during CO2 gasification at different temperatures.

mL/min (STP, volume hourly space velocity of 8.3 × 103 h−1). Therefore, all subsequent tests were performed within these limits. These limits ensured that the reaction rate and/or intraparticle diffusion resistance were the only controlling steps during combustion or gasification.19,20 During individual gasification tests, carbon conversion in the COC versus time could be used to obtain kinetic rate data. 2.4. Analytic Methods. In this study, COC was regarded as pure carbon. The small hydrogen content of the COC [hydrogen content of CLVR was 5.83 wt %, as determined by elementary analysis (EA 3000, HEKAtech)] was not considered in the data evaluation, i.e., kinetic data presented in this paper were all related to the carbon of the carbonaceous deposits. H2, CO, CH4, and CO2 were the principal products in the reaction. Gases were analyzed using a gas chromatograph (Agilent 6890N, N2 as carrier gas), which has two detectors. H2 composition was monitored using a thermal conductivity detector, and CO, CH4, and CO2 production were determined using a flame ionization detector. During isothermal measurements, initial and actual amounts of carbon on the catalyst were determined by burning the COC and determining the amount of CO 2 emitted during combustion. Coke was burned with a 150 mL/min pure oxygen flow for 2 min at 900 °C, and the resulting CO2 emission was quantified using an online infrared analyzer.

greater than 0.5 during CO2 gasification. Literature shows that coke not only distributes on the catalyst surface but also deposits in the catalyst pore.21−23 Knudson diffusion will be the dominant mode of diffusion when the ratio of the mean free path of the gas and pore size is greater than 10.24,25 The catalyst used in this study has an average pore size in the range 11−14 nm, and a pore size less than 10 nm accounts for ∼40 vol %. Within the conditions studied, the mean free path of CO2 is 160−180 nm. Clearly, Knudson diffusion gradually becomes serious as the reaction proceeds due to the remaining coke mostly deposited in pores with sizes well below the mean free path of CO2, so that the reaction rate decreases significantly. 3.2. Comparison of the Reactivities for COC during Steam and CO2 Gasification. The gasification reactivity is usually quantified by a reactivity index,26,27 Rs, defined as R s = 0.5/τ0.5

where τ0.5 represents the time required to reach 50% carbon conversion. This definition is commonly used to compare the gasification reactivities of different chars. Herein, this index is adopted to represent the reactivities of COC. The reactivity indices Rs for COC during steam, and CO2 gasification can be obtained from Figure 2. The relative reactivity ratio of steam to CO2 is defined as relative reactivity ratio = R H2O/ R CO2

3. RESULTS AND DISCUSSION 3.1. Effect of Temperature on COC during Steam and CO2 Gasification. The carbon conversion X is defined as the gasified carbon to the initial carbon in the coked-catalyst, as given below conversion X = (C i − Ca)/C i

(2)

(3)

where RH2 O represents the reactivity of steam and RCO 2 represents the reactivity of CO2. The reactivity indices Rs of COC gasification with steam and CO2 and the relative reactivity ratios of steam to CO2 at different temperatures are listed in Table 1.

(1)

Table 1. Comparison of the Reactivity Indices Rs for COC during H2O and CO2 Gasification at Different Temperatures

where Ci represents the initial carbon content of COC and Ca is the carbon content of COC at an intermediate gasification level. Panels a and b of Figure 2 illustrate the relationship of conversion (X) with time for COC during steam and CO2 gasification at 800−900 °C and 0.09 MPa of reactant partial pressure. It is seen that the carbon conversion increases with increasing reaction time, as expected. Under the same temperature, the rate of gasification of COC with steam is higher than that with CO2 at conversion lower than 0.5. Figure 2 also indicates that the gasification rate increases with increasing reaction temperature, however, a decrease of gasification rate as reaction proceeds for both reactions. The gasification rate decreases significantly when conversion is

temp, °C

RH2O (min−1)

RCO2 (min−1)

RH2O/RCO2

800 850 900

0.0145 0.0442 0.1199

0.0047 0.0093 0.0253

3.09 4.75 4.74

Many investigations on the gasification reactivities of char with steam and with CO2 have been conducted.26−30 However, studies on the comparison of reactivities and reaction mechanisms for steam gasification and CO2 gasification are limited. From Table 1, we can see that the gasification reactivities of COC with steam are 3−5 times higher than with CO2. The magnitude of these results is in good agreement with 16739



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that reported by other investigators.28,29 They found that the gasification of the reactivities of cokes (lignite and bituminous coal chars, commercial coke) gasified with steam is 2−5 times higher than with CO2. Ergun30 explains that the difference in rates for these two reactions is due to a greater number of active sites generated by steam than CO2. It seems that the reactivity of COC during steam or CO2 gasification is similar with that of some coal chars or cokes.

Figure 4. Conceptual process of coproduction syngas during regeneration of coked catalyst for upgrading heavy petroleum feeds.

syngas generator. A heavy feed is injected into the fluidized bed of hot catalyst in the riser. In this case, steam is used as the fluidizing medium. The gaseous and liquid products formed in the riser reactor are withdrawn at the top while coked catalysts are transported to the combustor where their temperature is increased by coke burning. The primary role of the combustor is partial decoking of coked catalysts to achieve a temperature, which is desirable for steam gasification in the syngas generator. Subsequently, hot semi-regenerated catalysts would be introduced into the syngas generator where they are contacted with steam, and then CO and H2 are produced. The catalyst is returned to the riser reactor, as a carrier of heat required for heavy feeds cracking, which exchanges heat, producing steam for the riser reactor and syngas generator. From the viewpoint of syngas generation, our process is very like the dual flexicoking process.31 The fundamental difference between the two processes is the “carrier”. For dual flexi-coking, the carrier is the solid coke particles, but for ours, the carriers are silica−alumina catalyst particles. The characteristics of the process in riser reactor have been studied in detail previously.11,12 More detailed reactions in combustor and syngas generator is given in following sections. 3.3.2. Reactions in Combustor. The energy required for the process is supplied by a partial burning of a coke layer deposited on the surface of catalyst particles in the combustor. To a certain extent, the amount of heat release depends on the nature of the coke. Because the coke on the catalyst is a complex carbonaceous material, a number of reactions happen in the combustor with different reaction heat. However, we will limit our study to C and H in this discussion because S and N are present in small amounts in the coke to be burned in the combustor. Thus, reactions 1−3 in Table 2 occur mainly in the combustor. Indeed, the Boudouard reaction (reaction 7) takes place even in the presence of oxygen.32 However, this reaction can be ignored under the conditions of this study. It should be noted that the products of carbon burning are CO and CO2; the heat release of the former is only 28% of the latter. The hydrogen content of COC is 5.83 wt % used in this experiment and coke can be represented as CH0.74. Thus, the heat release of burning hydrogen in CLVR is constant. Apparently, the amount

Figure 3. Reactivity indices Rs versus reaction temperature for COC during steam (■) and CO2 (●) gasification.

Reactivity index is plotted against the reaction temperature in Figure 3. If instead of using the rate constant, we use the reactivity index, by use of the Arrhenius law: k = A exp( −Ea /RT )

(4)

where A is the pre-exponential factor, Ea is the activation energy, R is the universal gas constant, and T is the reaction temperature. We can directly estimate the activation energy from the plots in Figure 3. The activation energy of CLVR gasification with steam and CO2 is 221.09 and 175.53 kJ/mol, respectively. It is also indicated that the gasification of CLVR depends strongly on the temperature, indicating chemicalreaction-rate control with the temperature range tested. 3.3. Conceptual Process Coproduction of Syngas during Regeneration of Coked-Catalyst Proposed. As is well-known, regenerating coked catalyst produces large amounts of CO2, which may cause global warming. Thus, it seems a good way that removing the coke on catalyst by CO2. However, one of the crucial problems is low reactivity of COC with CO2 at temperatures lower than 900 °C, at which destruction of structure or sintering of catalyst may not be avoided. And this temperature is much higher than a conventional FCC regeneration temperature of 650−750 °C. In addition, we have experimentally demonstrated the rate dramatically decreased when the conversion was higher than 0.5 during COC CO2 gasification. Thus, CO2 as gasifying agent does not seem suitable. Gasification reactivity of COC with steam is higher than with CO2. H2 and CO, as important raw materials for subsequent unit, could be produced simultaneously at relatively low temperatures during COC steam gasification. Therefore, we select steam as the gasifying agent for coproduction of syngas during regeneration of coked catalyst. 3.3.1. Description of the Conceptual Process. Simplified schematic of the process shown in Figure 4. The process consists of three parts, i.e., the riser reactor, combustor, and 16740



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consider as C. Related literature showed that the gasification products’ composition was influenced by temperature and pressure, which affect the extent of secondary reactions 5−7 (Table 2) during gasification and distribution of products.34 The effect of temperature and steam partial pressure on the distribution of products was shown in Tables 4 and 5,

Table 2. Main Reactions in Combustor and Syngas Generator Combustor C + (1/2)O2 → CO C + O2 → CO2 H + (1/4)O2 → (1/2)H2O Syngas Generator C + H2O → CO + H2 CO + H2O → CO2 + H2 CO + 3H2 → CH4 + H2O C + CO2 → 2CO

1. 2. 3. 4. 5. 6. 7.

ΔH = −110 kJ/mol ΔH = −393 kJ/mol ΔH = −121 kJ/mol ΔH ΔH ΔH ΔH

= = = =

Table 4. Semi-Regenerated Catalyst Gasification with Steam (90 kPa) at Different Temperatures

+131 kJ/mol −41 kJ/mol −206 kJ/mol +172 kJ/mol

carbon balance

of heat release decreased with the increasing CO/CO2 ratio in products, as we theoretically calculated in Figure 5. Many

temp, °C

CO

CO2

CH4

H2/CO molar ratio

750 800 850 900

18.5 18.7 19.9 27.3

7.5 6.1 4.2 6.0

1 1 1 1

1.79 1.73 1.77 1.85

Table 5. Effect of Steam Partial Pressure on Distribution of Products (vol %) during Semi-Regenerated Catalyst Gasification with Steam at 800 °C

studies show that the CO/CO2 ratio depends on temperature.32,33 Table 3 gives the CO/CO2 ratio at different Table 3. Effect of Temperature on the CO/CO2 Molar Ratio for COC Burning in Air CO/CO2

temp, °C

CO/CO2

600 630 660 690

1.05 1.21 1.40 1.60

720 750 780 810

1.18 0.75 0.26 0.15

H2

CO

CH4

CO2

55 71 83 90

56.00 55.75 55.26 55.64

30.65 31.53 31.76 32.17

1.13 1.32 1.30 1.42

12.22 11.40 11.68 10.77

respectively. By analyzing these data, we found that high temperature was in favor of yielding more syngas. Yet, the yield of syngas was 85 vol % in products at 750 °C. The effect of partial steam pressure on syngas yield was very small under the conditions of this study, yielding about 87 vol % syngas in products at 800 °C. According to above results, we calculated the equilibrium carbon content CLVR for temperatures for a combustor and syngas generator atmosphere. A controlled supply of air to the combustor will remove the coke layer and increase the temperature of particles to about 800 °C. All hydrogen in CLVR was consumed in the combustor. The hot particles were then transferred to the syngas generator where they were contacted with steam to produce CO and H2. Particles exiting the syngas generator were about 740 °C. Subsequently, these particles through heat exchange decreased to about 680 °C while they were transferred to the riser to supply the necessary catalyst for cracking reactions. In this case, about 60 and 40 wt % carbon on catalyst was removed in the combustor and syngas generator, respectively. 3.4. Kinetics of Regeneration. According to the process we have described in section 3.3.1, two steps were involved in regeneration of coked-catalyst: (a) CLVR burned in the combustor, wherein semi-regenerated catalysts (CSRC) were formed, and (b) CSRC entered the syngas generator gasification with steam. In the following sections, we will discuss those two steps separately. 3.4.1. Kinetics of COC Burning with Oxygen in the Combustor. Coke−O2 combustion is highly exothermic reaction, which causes the temperature of coked-catalyst particles to increase especially at higher oxygen concentrations, so that pore diffusion resistance may be significant and the reaction does not proceed under chemical reaction control. Therefore, in the present study, we adopt dilute oxygen in N2 (O2 2 vol %) to burn the coke deposited on the catalyst at temperatures of 650−800 °C. At lower oxygen concentrations,

Figure 5. Relationship between heat release and CO/CO2 in the reaction of COC with oxygen.

temp, °C

steam partial pressure, kPa

temperatures in products of CLVR burning with air. As can be seen in Table 3, the CO/CO2 ratio gently increases at temperatures of 600−690 °C and drastically decreases at temperatures of 690−810 °C. On the basis of value of CO/ CO2 at a specific temperature and the combustion heat of coke−O2 reactions, we can readily calculate that the amount of coke is consumed in the combustor, wherein a large amount of heat is generated. This heat increases the temperature of particles to the desired temperature. Then, these high temperature particles are transferred to the gasifier, wherein the remaining coke on catalyst gasification with steam can be well conducted. 3.3.3. Reactions in the Syngas Generator. Partially oxidized particles exiting the combustor into the syngas generator, where the remaining coke deposited on catalyst was used for the production of syngas according to reaction 4 in Table 2. Because the hydrogen−oxygen reaction rate was much faster than the carbon−oxygen reaction rate, the H in coke was consumed in the combustor and the remaining coke was 16741



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the temperature rise within the solid particles during CLVR combustion could be neglected and may decrease the oxidation rate to a level that would favor the reaction to occur throughout the interior surfaces of the CLVR particle, so that pore diffusion resistance is negligible and the reaction proceeds through chemical reaction control. Both combustion and gasification of coke deposited on catalyst is considered as a gas−solid reaction and generally follows first-order kinetics.17,35−37 Homogeneous and shrinking-core models, expressed by eqs 5 and 6, respectively, have been used widely for kinetic analysis of gas−solid reaction experimental data.20,26 dX = k(1 − X ) dt

or

dX = K (1 − X )2/3 dt

ln(1 − X ) = −kt or

1 − (1 − X )1/3 = Kt

Figure 7. Temperature dependence of ln k and ln K of COC combustion.

(5)

respectively, indicating that the reactions follow the Arrhenius law. The pre-exponential factor A and apparent activation energies Ea can be determined directly from the plots in Figure 7 as follows:

(6)

where X is the conversion and k and K are the kinetics reaction constant. Experimental results of CL‑VR during oxygen combustion can be described by both the homogeneous and shrinking core models. Both models respective experimental results at different temperatures are shown in Figure 6.

Homogeneous model for CLVR oxygen combustion: k = 2.34 × 105· exp( − 12496/T ) (A = 2.34 × 105 min−1 , Ea = 103.9 kJ/mol)

Shrinking-core model for CLVR oxygen combustion: K = 4.50 × 104 · exp( − 12176/T ) (A = 4.50 × 104 min−1 , Ea = 101.2 kJ/mol)

Interestingly, activation energies are very close for both models, and both have the same order of magnitudes but have different pre-exponential factors, suggesting that the diffusion effects are present but not significant. 3.4.2. Kinetics of Semi-Regenerated Catalyst Gasification with Steam in Syngas Generator. According to the results of CLVR combustion with oxygen, combined with the process analysis in section 3.3.3, we investigated the kinetics of semiregenerated catalyst (CSRC) gasification with steam in the syngas generator. By CLVR burning with limited oxygen at 750 °C, we prepared the CSRC with a carbon content of 0.95 wt %. Then, the CSRC gasification with steam experiment was carried out at 750−900 °C. Similarly, both the homogeneous model and shrinking-core model are used to describe the CSRC gasification with steam. Related results are shown in Figure 8. At temperatures higher than 800 °C, experimental data well match for both of the

Figure 6. Comparison of models and experimental data of COC burning with oxygen at 650−800 °C.

Both models appear to have fitted the experimental data well with conversions lower than about 0.7. The homogeneous model is found to be fairly better than the shrinking-core model, especially at lower temperatures due to slower reaction and faster oxygen diffusion through the pores of the particle at lower temperatures. However, both models have deviations as the conversion exceeded 0.7 at 650−800 °C. Because the catalyst used in this study is a porous material, the coke probably not only deposits on the catalyst surface but also deposits in catalyst pores during residua upgrading processing. Thus, as conversion increases, the coke on the catalyst surface is easy to react and be consumed, the remaining coke mainly deposits in pores, so that the internal diffusion resistance of the oxygen through the pores increases. Therefore, the controlling step may change during the reaction at higher conversion. Because the homogeneous model and shrinking-core model can be applied to COC gasification; the k and K in eq 5 and 6 can be estimated from the experimental data, respectively. The value of k and K for oxygen combustion of CLVR is plotted by the eq 4, as shown in Figure 7. The linear relationships between the k and K value and the reciprocal of reaction temperature are evident, with a correlation coefficient R2 of 0.9918 and 0.9887,

Figure 8. Comparison of models and experimental data of semiregenerated catalyst gasification with steam at 750−900 °C. 16742



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models. However, both models have deviations at a temperature of 750 °C. A possible reason for this result is that the reaction rate of carbon-gasification is slow at lower temperatures; thus the order of reaction deviates first, which is the prerequisite for the two models applied. Also, according to eqs 4, 5, and 6, values of k and K for steam gasification of CSRC are plotted, two straight lines with a correlation coefficient R2 of 0.9972 and 0.9971, respectively, as shown in Figure 9. And the pre-exponential factor A and apparent activation energies Ea can be determined directly from the plots in Figure 9 as follows:

mol for steam gasification and 175 kJ/mol for CO2 gasification. (2) A conceptual process of coproduction of syngas during regeneration of coked-catalyst by combining coke− oxygen combustion and coke−steam gasification is presented, where the coke deposited on catalyst was removed in a combustor and then in a syngas generator. The products yielded from the syngas generator containing H2 and CO at approximately 56 and 32 vol %, respectively. (3) Within the temperature range studied, both of the carbon-on-catalyst combustion with oxygen and gasification with steam are under chemical-reaction-rate control. The performances of CLVR combustion with oxygen and CSRC gasification with steam could be described by homogeneous model and shrinking-core model. For CLVR oxygen combustion, the value of activation energy is about 104 kJ/mol for homogeneous model and 101 kJ/mol for shrinking-core model. For CSRC steam gasification, the value of activation energy is about 226 kJ/mol for homogeneous model and 207 kJ/ mol for shrinking-core model.

■

AUTHOR INFORMATION

Corresponding Author

Figure 9. Arrhenius equations for homogeneous model and shrinking core model.

*G. Wang: e-mail, [email protected]; tel, 8610-8973-3085; fax, 8610-6972-4721.

Homogeneous model for CSRC steam gasification:

Notes

The authors declare no competing financial interest.

■

9

k = 1.86 × 10 · exp( − 27186/T )

ACKNOWLEDGMENTS The authors acknowledge the financial support provided by National Natural Science Foundation of China (21176252) and Key Technologies Research and Development Program of China (2012BAE05B02), the State Key Program of National Natural Science of China (21336011), and the Program for New Century Excellent Talents in University of China (NCET13-1029).

(A = 1.86 × 109 min−1 , Ea = 226.0 kJ/mol)

Shrinking-core model for CSRC steam gasification: K = 6.44 × 107 · exp( − 24897/T ) (A = 6.44 × 107 min−1 , Ea = 207.0 kJ/mol)

The reaction activation energy is 226.0 kJ/mol for the homogeneous model and 207.0 kJ/mol for the shrinking-core model. These results arapproach those of Corma et al.,17 where the activation energy for coked hydrotalcite FCC catalyst is 239 kJ/mol, showing that the regeneration rate is temperature dependent. The value of activation energy obtained by the homogeneous model is very close to that for the relationship of the reactivity index and temperature (see section 3.2). However, there is difference of about 20 kJ/mol for the activation energy and 2 orders of magnitude for the preexponential factor obtained with the two different models, implying that the internal diffusion effects are somewhat significant.

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REFERENCES

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4. CONCLUSIONS The following conclusions can be drawn from our present study: (1) The reactivity of coke deposited on catalyst with steam gasification is higher than that with CO2 gasification. The reactivity of the former is about 3−5 times higher than that of the latter at temperature range 800−900 °C. The values of activation energy were estimated by use of the relationship of reactivity index and temperature under these two different atmospheres. This value is 221 kJ/ 16743



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