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Figure 1. Simple process flow diagram of a GTL process [2].

Figure 1 Simple process flow diagram of a GTL process [2].

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The effect of main operating variables such as temperature, partial pressure of hydrogen, liquid and hydrogen flow rates on conversion of hydrogen and selectivity of cyclo-dodecene in consecutive hydrogenation of 1,5,9-cyclododecatriene to cyclo-dodecene was analyzed by Stuber and Delmas [16] in a multistage operation. Ngwenya et al. [17] examined the influ- ence of operating parameters on the performance of the cobalt catalyst-based FT reactor with different objective func- tions. Manenti et al. investigated possible design improve- ment in a series of fixed-bed reactors for methanol and dimethyl ether synthesis [18]. Hillestad applied the systema- tic staging method on a methanol reactor [19]. Staging will increase the production rate of methanol at the end of the reactor path. Staging of the FT reactor individually (single pass) was recently published for iron- and cobalt-based cata- lysts. A lumping technique was used to reduce the number Sectioning of chemical reactors into stages has been widely applied. Androulakis and Reyes investigated oxidative coupling of methane (OCM) on a multistage plug-flow reactor [10]. The effect of feed distribution in a packed-bed membrane reactor for methanol oxidative dehydrogenation was studied by Diakov and Varma [11]. Conversion of a staged slurry bubble column reactor for FITS was optimized by Maretto and Krishna [12]. Oxygen distribution and selective combustion of hydrogen during propane dehydrocyclodimerization reactions was per- formed by Waku et al. [13]. The effect of catalyst dilution and feed distribution of two case studies involving nitrobenzene hydrogenation and ethylene oxidation was investigated by Hwang and Smith [14]. Guillou et al. [15] studied the effect of distributed hydrogen in a microchannel FT reactor.
The effect of main operating variables such as temperature, partial pressure of hydrogen, liquid and hydrogen flow rates on conversion of hydrogen and selectivity of cyclo-dodecene in consecutive hydrogenation of 1,5,9-cyclododecatriene to cyclo-dodecene was analyzed by Stuber and Delmas [16] in a multistage operation. Ngwenya et al. [17] examined the influ- ence of operating parameters on the performance of the cobalt catalyst-based FT reactor with different objective func- tions. Manenti et al. investigated possible design improve- ment in a series of fixed-bed reactors for methanol and dimethyl ether synthesis [18]. Hillestad applied the systema- tic staging method on a methanol reactor [19]. Staging will increase the production rate of methanol at the end of the reactor path. Staging of the FT reactor individually (single pass) was recently published for iron- and cobalt-based cata- lysts. A lumping technique was used to reduce the number Sectioning of chemical reactors into stages has been widely applied. Androulakis and Reyes investigated oxidative coupling of methane (OCM) on a multistage plug-flow reactor [10]. The effect of feed distribution in a packed-bed membrane reactor for methanol oxidative dehydrogenation was studied by Diakov and Varma [11]. Conversion of a staged slurry bubble column reactor for FITS was optimized by Maretto and Krishna [12]. Oxygen distribution and selective combustion of hydrogen during propane dehydrocyclodimerization reactions was per- formed by Waku et al. [13]. The effect of catalyst dilution and feed distribution of two case studies involving nitrobenzene hydrogenation and ethylene oxidation was investigated by Hwang and Smith [14]. Guillou et al. [15] studied the effect of distributed hydrogen in a microchannel FT reactor.
Table 1. Kinetic and adsorption parameters in Eqs. (6) and (7) [9]. Iglesia’s reaction rate on a cobalt catalyst is valid at 473-483 K, 100-3000 kPa, H,/CO = 1-10, and CO conversions lower than 15%. It is assumed that this kinetic model is still valid for H;/CO <1 and CO conversion > 15%. It is known that applying Iglesia’s kinetic model outside the validity range is a possible source of error. This kinetic model is chosen because it takes into account a higher selectivity of methane. However, other available kinetic models could be applied, e.g., the kinetic model developed by Yates and Satterfield [24]. In their kinetic model, the consumption rate is measured in a well-mixed continuous-flow slurry reactor operating at T=220-240°C, P=0.5-1.5MPa, H,/CO of feed=1.5-3.5, conversion of H, and CO 6-68% and 11-73%, respectively. The mass fractions of FT products heavier than methane are assumed to follow an ideal Anderson-Schulz-Flory (ASF) dis- tribution where a is the chain growth probability. For cobalt catalysts, chain growth probabilities between 0.85 and 0.95 are reported [26].
Table 1. Kinetic and adsorption parameters in Eqs. (6) and (7) [9]. Iglesia’s reaction rate on a cobalt catalyst is valid at 473-483 K, 100-3000 kPa, H,/CO = 1-10, and CO conversions lower than 15%. It is assumed that this kinetic model is still valid for H;/CO <1 and CO conversion > 15%. It is known that applying Iglesia’s kinetic model outside the validity range is a possible source of error. This kinetic model is chosen because it takes into account a higher selectivity of methane. However, other available kinetic models could be applied, e.g., the kinetic model developed by Yates and Satterfield [24]. In their kinetic model, the consumption rate is measured in a well-mixed continuous-flow slurry reactor operating at T=220-240°C, P=0.5-1.5MPa, H,/CO of feed=1.5-3.5, conversion of H, and CO 6-68% and 11-73%, respectively. The mass fractions of FT products heavier than methane are assumed to follow an ideal Anderson-Schulz-Flory (ASF) dis- tribution where a is the chain growth probability. For cobalt catalysts, chain growth probabilities between 0.85 and 0.95 are reported [26].
Table 2. Optimal design of the GTL process with different numbers of stages of the FT reactors. The total productions of methane and water in the FT reac- tor are 140 and 6689 kmolh‘', respectively. The per pass con- versions of H, and CO in the FT reactor are 42.97 and 10.78 %, respectively. On the other hand, the total conversions of H and CO are 65.48 % and 23.21%, respectively. The amount of heat needed to be removed from the FT reactor is 1.0810’ kJ hr’. The flow rate of make-up natural gas as fuel to the fired heater is 1620 kmol h”!. In this case, 90 kmol h™ of CO, is emit- ted to the atmosphere via purge stream. gases, liquid hydrocarbon phase, and water. The light gases are sent to the second FT reactor. The effluent of the second FT reactor is sent to another three-phase separator to separate light gases, liquid hydrocarbon phase, and water. A part of the light gases from the second separator is purged and the rest is recycled back to the FT and syngas sections to increase the con- version of species.
Table 2. Optimal design of the GTL process with different numbers of stages of the FT reactors. The total productions of methane and water in the FT reac- tor are 140 and 6689 kmolh‘', respectively. The per pass con- versions of H, and CO in the FT reactor are 42.97 and 10.78 %, respectively. On the other hand, the total conversions of H and CO are 65.48 % and 23.21%, respectively. The amount of heat needed to be removed from the FT reactor is 1.0810’ kJ hr’. The flow rate of make-up natural gas as fuel to the fired heater is 1620 kmol h”!. In this case, 90 kmol h™ of CO, is emit- ted to the atmosphere via purge stream. gases, liquid hydrocarbon phase, and water. The light gases are sent to the second FT reactor. The effluent of the second FT reactor is sent to another three-phase separator to separate light gases, liquid hydrocarbon phase, and water. A part of the light gases from the second separator is purged and the rest is recycled back to the FT and syngas sections to increase the con- version of species.
Table 3. Production rate of hydrocarbons inside the FT reactor (single-stage reactor).
Table 3. Production rate of hydrocarbons inside the FT reactor (single-stage reactor).
Table 4. Carbon distribution of liquid hydrocarbon product.
Table 4. Carbon distribution of liquid hydrocarbon product.
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