Simulation and design study of an ethyl benzene chemical process

OPTIMIZATION STUDY OF FISCHER-TROPSCH GASTO-LIQUID (GTL) PROCESS USING ASPEN HYSYS MUHAMMAD FARRIS IDAHAM BIN ABDUL RASHID DIRECTOR: MANUEL JOSE LIS ARIAS 10th JUNE 2019 FINAL YEAR PROJECT (TFG) BACHELOR’S DEGREE IN CHEMICAL ENGINEERING Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS i CONTENTS ABSTRACT ............................................................................................................................................... iii ACKNOWLEDGEMENTS .......................................................................................................................... iv DECLARATION OF HONOUR .................................................................................................................... v LIST OF FIGURES ..................................................................................................................................... vi LIST OF TABLES ...................................................................................................................................... vii LIST OF ABBREVIATIONS ...................................................................................................................... viii 1. INTRODUCTION ............................................................................................................................... 1 1.1 INTRODUCTION TO FISCHER-TROPSCH TECHNOLOGY ........................................................... 1 1.2 PROCESS DESCRIPTION ........................................................................................................... 4 1.2.1 SYNTHESIS GAS PRODUCTION......................................................................................... 5 1.2.2 FISCHER-TROPSCH SYNTHESIS ........................................................................................ 7 1.3 SCOPE OF THE BACHELOR THESIS ......................................................................................... 12 2. OBJECTIVE ..................................................................................................................................... 13 3. LITERATURE REVIEW ..................................................................................................................... 14 3.1 SYNGAS PRODUCTION UNIT ................................................................................................. 14 3.2 FISHCER-TROPSCH SYNTHESIS .............................................................................................. 15 4. MODELLING IN ASPEN HYSYS ........................................................................................................ 17 4.1 AUTOTHERMAL REACTOR ..................................................................................................... 19 4.2 FISCHER-TROPSCH REACTOR ................................................................................................ 21 5. EXPERIMENTAL DEVELOPMENT .................................................................................................... 23 5.1 AUTOTHERMAL RECTOR OPTIMIZATION .............................................................................. 23 CASE 1: SYNGAS PRODUCTION UNIT WITH DIFFERENT REACTOR TYPES ..................................... 23 CASE 2: STAGING OF ATR WITH DOSSIFICATION .......................................................................... 25 5.2 FISCHER-TROPSCH REACTOR OPTIMIZATION ....................................................................... 26 CASE 3: INFLUENCE OF COBALT CATALYST CHARACTERISTIC ON PRODUCT DISTRIBUTION ........ 26 CASE 4: STAGING OF FTR WITH DOSSIFICATION........................................................................... 28 5.3 OVERALL PROCESS OPTIMIZATION ....................................................................................... 29 6. RESULTS AND DISCUSSIONS .......................................................................................................... 31 6.1 RESULT OF AUTOTHERMAL REACTOR OPTIMIZATION ......................................................... 31 6.1.1 DISCUSSION OF AUTOTHERMAL REACTOR OPTIMIZATION .......................................... 37 6.2 RESULT OF FISCHER-TROPSCH REACTOR OPTIMIZATION ..................................................... 39 6.2.1 DISCUSSION OF FISCHER-TROPSCH REACTOR OPTIMIZATION ..................................... 43 6.3 RESULT OF OVERALL PROCESS OPTIMIZATION ..................................................................... 45 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS ii 6.3.1 DISCUSSION OF OVERALL PROCESS OPTIMIZATION ..................................................... 47 7. CONCLUSION ................................................................................................................................. 49 BIBLIOGRAPHY ...................................................................................................................................... 50 ANNEX ................................................................................................................................................... 52 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS iii ABSTRACT Fischer-Tropsch (FT) technology is simply defined as the means to convert syngas to a wide range of hydrocarbon products including paraffins, olefins, oxygenates, etc. Interest in FT synthesis is increasing rapidly due to a number of factors including demand of cleaner fuel (low Sulphur, low aromatics), improvement in FT technology and the need to monetize stranded natural gas resources. Hence, there is a strong incentive to further optimize the reactions in a GTL plant so that the overall cost can be reduced. In this study, the overall GTL process was simulated using Aspen HYSYS v8.6. The impact of different parameters on the performance of syngas unit, FTS unit and the overall process were studied. The main objective is to optimize all the main reactions involved in GTL process as well as to increase fraction of desired product. To achieve this, all the important reactions were modelled using kinetic approach obtained from literature and kinetic reactors were used. The Auto Thermal Reactor (ATR) was modelled as CSTR and PFR in HYSYS with kinetic data given by Xu,Froment et al. and D.L Hoang et al. [11], [22]. The main objective function was to increase conversion of reaction and to produce low H2:CO ratio in syngas. Three configurations were proposed to compare the contact model of the ATR using different reactors. CSTR was observed to produce syngas with suitable H2:CO ratio for FTS. Staging of ATR with dossification of feed between stages was simulated however it did not show any improvement on conversion or H2:CO ratio. The FT reactor with cobalt based catalyst was modelled as PFR with kinetic data from Iglesia et al.[27] A few parameters were varied to observe the effect on product distribution and overall conversion. In case 3, the effect of varying void fraction in FTR was studied and the optimal value was recorded. Case 4 explored the effect of multiple stage FTR configuration. Staging of FTR recorded a higher overall conversion and product distribution also improved. An optimal FTR configuration was proposed and showed an even higher conversion compared to the multi stage FTR and an improved product distribution. Some other process parameters were also investigated such as the inlet temperature of ATR and steam:carbon ratio in pre-reformer. Optimization of both parameters showed improvement in the syngas produced. Lower H2:CO ratio of syngas was observed by increasing inlet temperature of ATR and reducing steam:carbon ratio. The optimal GTL plant configuration was proposed at the end of this study. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS iv ACKNOWLEDGEMENTS This work has been carried out at the Department of Chemical Engineering at The Universitat Politecnica de Catalunya (UPC) - BarcelonaTECH in Terrassa, Barcelona. I would like to take this opportunity to thank and give acknowledgements to those who have helped me complete this bachelor’s thesis be it directly or indirectly. First and foremost, I wish to express my deepest gratitude to my supervisor, Professor Manuel Jose Lis Arias for his invaluable guidance, support and encouragement during the course of this semester. I am truly grateful that he always makes time out of his busy schedule to attend to my questions and doubts. I have really enjoyed learning and working with such a complete academician like Professor Manuel and I am glad that I have learned a lot from him. I would also like to thank all the lecturers and professors who have taught me during my four years in UPC. I have learned and grown so much since I first step into the doors of this university. The challenges I face during my bachelor’s degree have certainly helped me to be stronger and wiser in life. It has been a wonderful experience completing my bachelor’s degree in such a beautiful city like Barcelona. Special thanks to all my friends here in Spain as well as back home in Malaysia. Their encouragement and support have helped me in many different ways to complete my studies. I shared so many good memories and laughter with my friends and they made my journey here in Spain easier and more enjoyable. Many thanks to the Government of Malaysia and MARA for funding me throughout my studies in Barcelona as well as the 2 years in Malaysia where I did my A-levels. It would not have been possible without their financial support and it has surely helped me to solely focus on my studies. Finally, this work would not have been possible without the love and support from my parents and family in Kuala Lumpur, Malaysia. Their encouragement, trust, motivation and invaluable support have made me who I am today. I will always treasure and appreciate everything they have done for me. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS vi LIST OF FIGURES Figure 1-1: A simplified process scheme of FTS [3] ................................................................................ 2 Figure 1-2: GTL process block diagram [2] .............................................................................................. 3 Figure 1-3: Process flow diagram of GTL process [6] .............................................................................. 4 Figure 1-4: Variation of product weight fraction with chain growth probability [1] .............................. 9 Figure 1-5: Multitubular fixed-bed reactor used in FTS [18] ................................................................ 11 Figure 3-1: syngas production unit process scheme [19] ..................................................................... 14 Figure 3-2: Chain building of FT products [1] ........................................................................................ 16 Figure 5-1: Simulation of ATR modelled as CSTR .................................................................................. 23 Figure 5-2: Simualtion of ATR modelled as isothermal PFR .................................................................. 23 Figure 5-3: Simualtion of ATR modelled as adiabatic PFR ................................................................... 24 Figure 5-4: Simulation of multiple stage ATR ....................................................................................... 25 Figure 5-5: HYSYS catalyst parameeters ............................................................................................... 26 Figure 5-6: Simulation of single stage FTR ............................................................................................ 27 Figure 5-7: Simulation of multi stage FTR ............................................................................................. 28 Figure 5-8: Optimal multi stage FTR configuration ............................................................................... 29 Figure 5-9: Final GTL process configuration .......................................................................................... 30 Figure 6-1: variation of molar fraction of desired product with catalyst void fraction in FTR ............. 40 Figure 6-2: Variation of inlet temperature of ATR with H2:CO ratio in syngas ..................................... 45 Figure 6-3: Effect of steam on H2:CO ratio in syngas ............................................................................ 46 Figure 6-4: Optimal GTL plant configuration ........................................................................................ 48 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS vii LIST OF TABLES Table 1-1: Effect of changing of parameters on chain growth probability, olefin/parafin ratio, carbon deposition and methane selectivity [3] .................................................................................................. 9 Table 1-2: Ideal characteristic of FTR [18] ............................................................................................ 11 Table 3-1: Application of different catalysts used in FTS and their corressponding value of  [2] ..... 16 Table 4-1: Molar composition of natural gas [1],[2] ............................................................................. 17 Table 4-2: Chemical reactions in pre-reformer and corresponding enthalypy values [6] .................... 18 Table 4-3: Feed process conditions [9],[2] ............................................................................................ 18 Table 4-4: Chemical reactions in ATR [6],[11] ....................................................................................... 19 Table 4-5: Kinetic data of reactions in ATR [11],[22] ............................................................................ 20 Table 4-6: Chemical reactions in FTR [6] ............................................................................................... 21 Table 4-7: Kinetic data of reactions in FTR [6] ...................................................................................... 22 Table 6-1: Molar fraction data of configuration 1 ............................................................................... 31 Table 6-2: Process condition data of configuration 1 ........................................................................... 32 Table 6-3: ATR (CSTR) data.................................................................................................................... 32 Table 6-4: Molar fraction data of configuration 2 ............................................................................... 33 Table 6-5: ATR (Isothermal PFR) data ................................................................................................... 33 Table 6-6 Molar fraction data of configuration 3 ................................................................................ 34 Table 6-7: ATR (Adiabatic PFR) data ..................................................................................................... 34 Table 6-8: Molar fraction data of multi stage ATR ............................................................................... 35 Table 6-9: Multi stage ATR data ............................................................................................................ 35 Table 6-10: Process condition data of multi stage ATR simulation ...................................................... 36 Table 6-11: Summary of case study 1 .................................................................................................. 37 Table 6-12: Summary of case study 2 ................................................................................................... 38 Table 6-13: Variation of molar fraction of desired product with catalyst void fraction in FTR ............ 39 Table 6-14: single stage FTR data .......................................................................................................... 40 Table 6-15: Molar fraction data of multiple stage FTR simulation ....................................................... 41 Table 6-16: Reactor data of multiple stage FTR .................................................................................... 42 Table 6-17: Process condition of multiple stage FTr simulation ........................................................... 42 Table 6-18: Summary of multi stage FTR data ...................................................................................... 44 Table 6-19: Reactor data of optimal FTR configuration ........................................................................ 45 Table 6-20: Molar fraction of optimal ATR configuration ..................................................................... 46 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS viii LIST OF ABBREVIATIONS FT - Fischer-Tropsch FTS - Fischer-Tropsh Synthesis FTR - Fischer-Tropsch Reactor GTL - Gas-to-Liquid ATR - Auto Thermal Reactor HTFT - High Temperature Fischer-Tropsch LTFT - Low Temperature Fischer-Tropsch CSTR - Continous Stirred Tank Reactor PFR - Plug Flow Reactor LNG - Liquified Natural Gas WGS - Water Gas Shift  - Chain Growth Pobability  - Weight Fraction U - Hydrogen Usage Ratio  - Catalyst Diffusional Coefficient  - Coefficient of First-order rate constant and effective diffusivity of -olefin  - Coefficient of structural catalyst properties Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS ix Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 7 1.2.2 FISCHER-TROPSCH SYNTHESIS The Fishcer-Tropsch process is the heart of the GTL process where syngas is converted to linear gaseous, liquid and solid hydrocarbon chains with the use of appropriate catalyst. The FTS is in principle a carbon chain building process, where CH2 groups are attached to the carbon chain thus forming a product with long hydrocarbon chain.[10] Which reactions exactly takes place and how, is a matter of controversy, as it has been in the last century since 1930’s.[12] FT synthesis produces a wide range of products as mentioned before but the two main products are alkanes and alkenes. These reactions are highly exothermic, and to avoid an increase in temperature, which results in lighter hydrocarbons, it is important to have sufficient cooling, to secure stable reaction conditions. The total heat of reaction amounts to 25 % of the heat of combustion of the synthesis gas, and lays thereby a theoretical limit on the maximal efficiency of the FT process.[10] This reaction involves a catalyst to be used in different types of reactor, mostly an iron or cobalt based catalyst where the reaction takes place. FTS is divided into two processes which are low or high temperature process (LTFT, HTFT) with temperatures ranging between 200-240 °C for LTFT and 300350 °C for HTFT.[12] If the desired products are shorter hydrocarbon chain for example, naphtha and gasoline, the most used reactor type is the gas-fluidized bed reactor operating in the HTFT temperature range with typically an iron catalyst.[2] In high temperature FTS, fixed bed reactors are not suitable to be used. This is because of the fact that carbon deposition occurs at higher temperatures which will block the reactor tubes and may poison the cobalt catalyst.[1] On the other hand, if heavier hydrocarbons for example diesel and wax are desired, multi-tubular fixed bed reactor or slurry bubble column reactor types are more preferable which operates in the LTFT temperature range.[2] In low temperature FTS, both precipitated iron catalysts or supported cobalt catalysts may be used. Catalyst poisoning is still a risk in this process though lower compared to HTFT. This study will focus more on LTFT process and will not take into account catalyst poisoning. As stated above, the two most used catalyst in FTS are cobalt and iron-based catalyst. Iron is cheap, but cobalt has the advantage of higher activity and longer life, though it is many times more expensive than iron catalyst.[12] On the other hand, water gas shift (WGS) reaction is active for iron catalyst but negligible for cobalt catalysts, leading to improved hydrocarbon yield.[13] Therefore, we can say that cobalt is a more suitable catalyst to be used in FTS. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 8 The two main reactions involved in the FTS are as follow:[6] CO+ U H2 → i=1 20 (ASF Coef.)iCiH2i+2+(ASF Coef.)30C30H62+ H2O (1) CO+3H2  CH4+ H2O (2) For the FT reaction, only paraffin products are considered in this work. Since the main product that we aim to maximize is diesel, hence the value of  is taken as 0.9. Figure 1-4 below justifies our selection of value of  = 0.9. This gives a hydrogen usage ratio,U of 2.1 given by equation 3. The stoichiometric coefficients are then calculated using the equations as outlined in a paper by Hillestad.[14] 𝑈=3 −  (3) 𝑟𝐹𝑇 =(1−  )2  𝑖−1 for Ci ,i=1,…..20 𝑟𝐹𝑇 =(1−  )  20 for C21→, Carbon distribution in the Fischer-Tropsch synthesis is determined by the probability of chain growth on the catalyst, also called the -value. The product distribution is often explained by a statistical distribution called Anderson, Schultz, and Flory (ASF), given by:[2]  𝑖=𝑖(1−  )2  𝑖−1 where i is the number of carbon atoms, and i is the weight fraction of a component with length i. Probability for the chain to terminate is explained by 1-. A graph of weight fraction of product against chain growth probability is depicted in figure 1-4:[1] Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 9 Figure 1-4: Variation of product weight fraction with chain growth probability [1] As the graph shows, the optimum chain growth probability, which corresponds to our desired product, diesel (C9− C25) is   0.9. For cobalt catalyst, Song et al.[15] found the dependency of chain growth probability as a function of H2:CO ratio and temperature. Low ratio H2/CO will increase chain growth probability, and therefore the proportion of produced heavier hydrocarbons is increased. [16] The effect of temperature as explained in the previous section regarding the LTFT and HTFT also affects the chain growth probability as higher temperature causes lower chain growth probability. Other parameters that could influence chain growth probability can be summarized in table 1-1.[3] We will however not consider the formation of olefins in this study as the quantity is very small compared to the paraffin products at a specific operating condition. Table 1-1: Effect of changing of parameters on chain growth probability, olefin/parafin ratio, carbon deposition and methane selectivity [3] Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 10 There are currently four main types of reactor used in the FTS:[13], [17] • Tubular fixed bed reactor • Slurry phase reactor • Circulating fluidized bed reactor • Fluidized bed reactor Of the four mentioned reactors, only two of them are commercially used today, the fixed bed reactor and slurry bubble column reactor. The world’s largest GTL plant producing 140,000 bbl/day of petroleum liquids uses a fixed-bed reactor technology while a slurry bubble column reactor maximum production capacity as of today is about 34,000 bbl/day. Selection of type of reactor used is essential for a large-scale plant. Multitubular reactors are widely applied to carry out exothermic reactions. They are easy to handle and to design because the parallel tubes behave very similarly. They are also easier to be scaled-up which is essential for large scale GTL plants and also helps to make the process more economical. Furthermore, there is no need for any equipment to separate the heavy wax from the catalyst particles because the liquid trickles down the catalyst. However, multitubular reactors also exhibit disadvantages such as high pressure drop and high capital cost.[18] The second reactor type which is industrially used for FTS is the bubble column reactor with suspended catalyst. In slurry bubble columns, fine catalyst powders with dimensions of 10 to 200 µm are used.[18] Thus, the influence of internal mass transfer resistances are negligible and optimal activity and selectivity can be achieved. Internals assure efficient heat removal from the reactor that allows for a nearly isothermal operation. Even though the achievable catalyst fraction of up to 25 vol.-% is lower than in fixed bed reactors (up to 60 vol.- %), the reactor productivity of a slurry bubble column should be higher due to the enhanced catalyst utilization and the higher average reactor temperature. [18] Nevertheless, two aspects decelerated the commercialization of slurry bubble columns for FTS. Both separation of the solid catalyst from the liquid products and the scale-up of these reactors are major challenges for industrial use of slurry bubble columns.[18] Besides that, in a case of catalyst poisoning, all catalyst in a bubble column reactor would be affected compared to a fixed-bed reactor where just a portion of catalyst would be affected. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 11 The advantages and disadvantages of both reactors mentioned can be summarized in table 1-2[18] It can be concluded that an ideal reactor concept has not been developed yet but it should have the following characteristics:[18] • Fixed bed catalyst • High catalyst efficiency due to short diffusion distances • Highly efficient gas-liquid mass transfer • Isothermal operation at highest possible temperature Table 1-2: Ideal characteristic of FTR [18] Based on the reasons and discussions mentioned earlier, it has been suggested that this study will model a fixed-bed reactor for the FTS. Figure 1-5 depicts a multitubular fixed-bed reactor used in FTS process.[18] Figure 1-5: Multitubular fixed-bed reactor used in FTS [18] Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 12 1.3 SCOPE OF THE BACHELOR THESIS A study of the GTL process covers a very wide area involving many processes and chemical reactions. We can see that there are many variables that need to be considered when aiming to optimize the GTL-FT process. In this section, the main area of study and key variables will be specified. To be able to study and optimize the GTL process, we will model and simulate the whole process using Aspen HYSYS v8.6 simulation software. The main area of this study will be on optimization of chemical reactions in syngas production unit and FTS while also taking into account optimal condition of the process. To achieve this, we will adopt a kinetic approach to model all the main reactions in the GTL process by using Aspen HYSYS. This software will enable us to predict the behavior of the process and analyze the response of the system by changing key variables. We should also be able to understand better the chemical reactions involved in the GTL process and this will allow us to optimize the reactions and therefore increase the efficiency of the process. As stated before, the syngas production unit and FTS will be considered while leaving out product upgrading unit in this study. For the syngas production unit, the ATR will be studied further since it is widely used in most commercial GTL plants today. For the FT reaction, the LTFT process will be studied in depth because it produces heavier hydrocarbon chains which are valuable and easily cracked to lighter ones if required. Since products of FTS consist of various hydrocarbons, only the main product which is paraffin will be considered. The reactor of FTS that we will focus on is the fixed-bed catalytic reactor with the use of cobalt catalyst while not considering risk of catalyst deactivation. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 13 2. OBJECTIVE Based on the challenges faced by the industry to increase the efficiency and finding the optimal condition of the Fischer-Tropsch (FT) synthesis, this project is aimed to achieve the following objectives: I. To model all the main reactions of the GTL process by using kinetic approach II. To explore the possibility of staging of reactors in syngas production unit by using a kinetic approach III. To maximize the production of longer hydrocarbon chains (C13− C19) by optimizing the reactions in both autothermal and FT reactor In the following part of this report, a collection of works done by other researchers regarding the optimization of GTL process will be further explained and discussed. This will help us to have a better understanding of the challenges of this process and the ways to overcome them in order to achieve the stated objectives. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 14 3. LITERATURE REVIEW 3.1 SYNGAS PRODUCTION UNIT The syngas production unit suggested by Bakkerud [19] is shown in figure 3-1. The syngas production process includes a pre-reformer, fired heater and ATR. Figure 3-1: syngas production unit process scheme [19] Typically, ATR is used to produce syngas in a GTL process as suggested by many researches in their work because it is regarded as the most economical and suitable method to produce syngas.[2], [4]– [8], [10], [20]. However, Rafiee et al. [1] has studied the combination of ATR and steam methane reforming to produce syngas called the combined reformer. He found out that by installing a steam methane reformer together with an ATR in the syngas production unit, production rate of syngas can be increased by 25.3% and total oxygen consumption will be reduced[1]. In Haldor Topsøe design (Figure 3-1), the inlet temperature to the ATR is 650°C and it is a tradeoff between available piping material and process economic considerations. The higher temperature the more oxygen is saved. The steam:carbon (S:C) ratio can affect the syngas process. If a low S:C ratio Is used, oxygen consumption in the ATR at a specific reaction temperature is also lower. But a low S:C ratio risks formation of soot. Commercial use has been proven by Haldor Topsøe with a ratio of 0.6. Even lower ratios, down to 0.2 have been reported [7] and commercialization of an ATR technology with S:C = 0.4 is expected within the near future.[1] The H2:CO ratio depends on the inlet composition, S:C, O2:C, inlet temperature, etc[1], [21]. A few studies have shown that by adding or recycling rich CO2 gas to the ATR, the equilibrium of the WGS can be shifted and lower H2:CO ratio in syngas can be produced.[1], [20] Higher outlet temperature of Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 15 ATR can also reduce H2:CO ratio but typically it is limited to 1030 °C due to material constraints and stability of catalyst in reactor.[1] After scrutinizing many academic journals and theses, almost all researches modeled the ATR as a conversion or Gibbs reactor in their simulation study by assuming chemical equilibrium at the outlet of ATR. [2], [4]–[8], [10], [20] Within this study, the ATR was modeled as a kinetic reactor in Aspen HYSYS with the kinetic data obtained from Xu,Froment et al. and D.L Hoang et al.[11], [22]. The oxygento-carbon ratio in the ATR was adjusted to reach the temperature of 1030 °C at the outlet of the reactor as proposed by Magne Hillestad et al.[23] The effect of S:C ratio was studied further and will be optimized to produce low H2:CO ratio in syngas. 3.2 FISHCER-TROPSCH SYNTHESIS Most work done by researchers studying the FTS has either modeled the FT reactor as slurry bubble column reactor or multitubular fixed bed reactor as they are the most widely used commercially today. The majority of research also focus on LTFT process instead of HTFT because of the more valuable products in LTFT. A number of ways have been developed to optimize the FTS to increase its efficiency. Works by Rafiee et al.[1], K.Moen et al.[2], C.Maretto et al.[24] and Magne Hillestad et al.[25] have all suggested that staging of the FTR could increase the conversion and improve efficiency. This method was first proposed by M.Hillestad[26] with kinetic data from Iglesia et al. based on cobalt catalyst.[27] They argued that with staging of FTR, mass fraction of desired product and CO conversion improves significantly. Rafiee [1]suggested that optimal mixing structure of a 2-stage FTS is CSTR for the first stage and PFR for the second. He stated that the CSTR will level out the temperature peak before entering the second stage. K.Moen [2] suggested that extra H2 feed to be injected between stages of the FTR so that the ratio of H2:CO remains constant between each stage as he noted that H2 is used up more than its stoichiometric coefficient compared to CO. He also argued that by removing water after every stage will increase the conversion in the FTR even higher, stating the fact that the residence time will increase as the feed is not diluted with water. However, he noted that this involves a higher capital cost as the need for more 3-phase separators, coolers and heaters and also causing the increase in utility cost. K.Moen [2] concluded that the best configuration of the FTS would be a 3 stage FTR and some of the steam is replaced by CO2 which would result in higher energy and carbon efficiency. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 16 The product distribution of the FTS synthesis can be explained by the ASF distribution as suggested by many researchers[1], [2], [6]. As the FTS in principle a chain building process, the chain growth probability according to ASF distribution theory can be called  and the chain termination probability is called 1-. The figure 3-2 depicts the chain building process.[1] Figure 3-2: Chain building of FT products [1]  depends on a number of things including temperature and type of catalyst used. Table 3-1 shows the typical value of  for different catalysts and products. [2] Table 3-1: Application of different catalysts used in FTS and their corressponding value of  [2] In this study, the FTS synthesis was modeled as a PFR in Aspen HYSYS as the idea is to simulate it as a multitubular fixed bed reactor. K.Moen[2] and Otaraku et al.[6] also did the same by modelling the FTR as PFR in their simulation study. However, O.Kristian [8] modeled the FTR as a CSTR citing that he models the reactor as a slurry bubble column reactor. Cobalt catalyst used will be modeled in the simulation and the kinetics data is given by Iglesia et al.[27] Staging of FTR as proposed by Hillestad[26] and studied by Rafiee[1] was investigated and optimized while finding the optimal configuration as suggested by K.Moen[2] in his study. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 23 5. EXPERIMENTAL DEVELOPMENT 5.1 AUTOTHERMAL RECTOR OPTIMIZATION CASE 1: SYNGAS PRODUCTION UNIT WITH DIFFERENT REACTOR TYPES In case 1, the effect of different reactor types in syngas production unit was studied and simulated in Aspen HYSYS. The main objective function to be optimized is the H2:CO ratio which should be around 2. CSTR and PFR were used to model the ATR in the syngas production unit. Firstly, the syngas production unit was simulated using CSTR and the results were recorded. Then, the CSTR was replaced by the PFR and simulated again. The performance and results of both of the reactors were compared to each other. This was done to study the optimal contact model of the ATR to produce syngas with H2:CO2. The flowsheet of the simulated syngas process configuration in HYSYS with CSTR and PFR is shown in figure 5-1 and figure 5-2 respectively. Configuration 1 Figure 5-1: Simulation of ATR modelled as CSTR Configuration 2 Figure 5-2: Simualtion of ATR modelled as isothermal PFR To FTR Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 24 Both configurations in figure 5-1 and 5-2 were simulated with the same kinetic reaction set and kinetic data. The operating condition was also set to be constant to facilitate comparison of the reactors. All reaction details and operating condition can be found in section 4.1. Another configuration was also simulated as shown in figure 5-3. This configuration is similar to the figure 5-2 except that the ATR was modelled as an adiabatic reactor by adjusting the oxygen flow rate using the adjust block so that the outlet temperature of the ATR remains constant at 1030 °C. Configuration 3 were compared to configuration 2 to study the effect of adiabatic and isothermal PFR on overall conversion. All three configurations were simulated with the same reaction set and process condition to facilitate comparisons. Configuration 3 Figure 5-3: Simualtion of ATR modelled as adiabatic PFR The pressure drop in the heaters and coolers were set to be 100 kPa whereas in the ATR it was set to 200 kPa. The results of all 3 configurations were recorded and compared which can be seen in the next chapter. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 25 CASE 2: STAGING OF ATR WITH DOSSIFICATION In this case, the effect of staging of ATR in the syngas production unit was investigated. In this configuration, multiple ATRs were placed in series and pure oxygen was injected into the 2nd and subsequent reactors. This was done to ensure a constant CH4:O2 ratio in all the ATRs since it was observed that the outlet of each ATR has a higher CH4:O2 ratio compared to the inlet. The objective function was to obtain syngas with H2:CO2 and to increase conversion. The flowsheet of case 2 configuration can be seen in figure 5-4. In case 2, the steam:carbon ratio was reduced to 0.5 based on the result of case 1 so that it produces syngas with lower H2:CO ratio. As shown in figure 5-4 above, the case 2 configuration was simulated with multiple PFRs modelled as ATR with multiple stages. The total volume of the reactors was kept constant in all configurations. This is to ensure that the volume of reactor does not affect the conversion. All three PFRs was attached with the same reaction set and the operation conditions was also kept constant in all configurations. Between each PFR, a cooler was placed to cool down the reactant to the reaction temperature which is 675 °C. In the first ATR, the reactor was modelled as adiabatic reactor similar to the configuration shown in figure 5-3. In the subsequent reactors, they were modelled as isothermal reactors with the same CH4:O2 ratio used in the first reactor. This was done because HYSYS takes a long time to simulate all three reactors as adiabatic reactors with adjust blocks. This however would not affect our result since in each reactor, the inlet composition of reactants remains constant with the help of oxygen dossification. The result of case 2 configuration was recorded and will be discussed in the next chapter. Figure 5-4: Simulation of multiple stage ATR Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 26 5.2 FISCHER-TROPSCH REACTOR OPTIMIZATION CASE 3: INFLUENCE OF COBALT CATALYST CHARACTERISTIC ON PRODUCT DISTRIBUTION The next case study that was conducted involves the FTR which produces hydrocarbon products. In case 3, the influence of variation of catalyst characteristic in FTR on product distribution was investigated. Among the catalyst characteristics that were studied are: • Void fraction • Particle sphericity • Particle diameter • Solid density This was done using Aspen HYSYS simulation software. HYSYS permits the variation of catalyst parameters under the reactions tab as shown in figure 5-5 below. The void fraction parameter can be changed in the rating tab of the PFR. Figure 5-5: HYSYS catalyst parameeters Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 27 The objective of this case study is to maximize the production of middle distillate as it is our desired product while also taking into account overall conversion. The hydrocarbon chain the we aim to maximize is in the range of C11− C19. Only one parameter was changed at a time to observe the effect on product distribution. Figure 5-6: Simulation of single stage FTR The flowsheet of this case study simulated in Aspen HYSYS can be seen in figure 5-6. The effect of product distribution was observed by comparing the composition data of the “crude product” stream as shown in the figure above. A few graphs were plotted showing the variation of product distribution with respect to change in catalyst parameter. The result of this case study will be shown in the next chapter. As seen in the figure above, a single stage FTR was simulated. The result of this case study will be compared with the next case study which involves multiple stage FTR. Total volume of reactor will be kept constant and optimum void fraction will be used in all FTR. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 28 CASE 4: STAGING OF FTR WITH DOSSIFICATION The next case study that was conducted is the effect of staging of FTR to increase conversion and improve product distribution. The configuration proposed for this case study is somewhat similar to case 2: Staging of ATR but with multiple PFRs arranged in series and dossified with pure hydrogen. The flowsheet as simulated in HYSYS is shown in figure 5-7. Figure 5-7: Simulation of multi stage FTR The syngas produced from the ATR passed through a heater to heat it up to 210 °C before being fed to the FTR. The FTR was modelled as an isothermal reactor as the outlet temperature was set to 220°C. To model this in HYSYS, an energy stream was attached to the PFR. The outlet of the first ATR was then cooled to 210°C by a cooler and then dossified with pure H2 stream in a stream mixer. This was done to ensure constant H2:CO ratio of reactants at the entry of each reactor since it was observed that H2:CO ratio decreases at the exit of each reactor. This shows that the consumption of H2 is slightly higher than its stoichiometric coefficient. At the exit of the last FTR, the exit stream was cooled to 30°C to condensed the vapor and was passed through a 3-phase separator. The 3-phase separator separated the aqueous, liquid and gas streams from each other. The aqueous stream which consisted of H20 was discharged and the unreacted gas stream was purged. The gas stream could also be recycled into the system but it was not considered in this study. The liquid phase which consisted of the main product was sent to the product upgrading unit. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 29 5.3 OVERALL PROCESS OPTIMIZATION After studying all the case studies proposed as mentioned before, an optimum configuration was proposed. The optimal configuration has taken into account the optimization of both ATR and FTR which includes staging of reactors, optimal reactor type and optimal catalyst parameters. Furthermore, a few optimizations of operation condition have also been proposed. Among the proposed optimization of operation condition are: I. Variation of inlet temperature of ATR to improve H2:CO ratio II. Separation product and by-product after each stage of FTR III. Single stage ATR IV. Optimum Steam:Carbon ratio For the first proposed optimization, the inlet temperature was varied by changing the inlet temperature of pure oxygen. The main parameter to be optimized is the H2:CO ratio. The objective is to obtain H2:CO2 by varying the inlet temperature. A graph was plotted to show the effect of inlet temperature in ATR on the H2:CO ratio of syngas. The result is shown in the next chapter. The next proposed optimization is separating the product and by-product between each stage of the FTR. The objective function is to increase conversion and improve product distribution so that more desired product is formed. The process flow diagram as simulated in HYSYS can be seen in figure 5-8. Figure 5-8: Optimal multi stage FTR configuration Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 30 The exit of the first FTR is cooled before entering a 3-phase separator. Here, the separation of unconverted syngas from crude product and water takes place. Only the unconverted gas is passed through the second FTR for the subsequent FT reaction. This configuration should increase the residence time and decrease duty of heater. However, this configuration will also cause higher capital cost as extra 3-phase separators and heaters are needed. The effect of separation of product, byproduct from unreacted syngas on overall conversion was studied and the results were recorded. The final configuration which was determined as the optimal configuration is shown in figure 5-9 below. The results of all the proposed configurations and case studies are presented in the next chapter. A thorough analysis was done and the discussions are presented in chapter 6 of this thesis which can be found in the next page. Figure 5-9: Final GTL process configuration Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 31 6. RESULTS AND DISCUSSIONS 6.1 RESULT OF AUTOTHERMAL REACTOR OPTIMIZATION Case 1: Syngas production unit with different reactor types The case 1 configuration was simulated in Aspen HYSYS as described in section 5.1. As mentioned in said section, there are 3 different configurations which was investigated. Firstly, we will compare configuration 1 to configuration 2. This comparison is done to observe the effect of different reactor modelled as the ATR on H2:CO ratio. In this case, the ATR in configuration 1 was modelled as CSTR while in configuration 2 it was modelled as PFR. Both reactors have the same volume. We will now observe the result of the simulation of these two configurations. Among the parameters to be considered are: • H2:CO ratio • Overall conversion Configuration 1 Table 6-1 shows mole fraction composition of all relevant streams which involve the ATR. Table 6-1: Molar fraction data of configuration 1 NG TO PREREFORMER DUMMY1 OUT PREREFORMER STEAM IN ATR Oxygen DUMMY2 OUT ATR Comp Mole Frac (H2O) 0 0 0.333 0.333 1 0.333 0 0.207 0.207 Comp Mole Frac (Oxygen) 0 0 0.000 0.000 0 0.000 1 0.006 0.006 Comp Mole Frac (Methane) 0.95 0.95 0.611 0.611 0 0.611 0 0.000 0.000 Comp Mole Frac (Ethane) 0.02 0.02 0.000 0.000 0 0.000 0 0.000 0.000 Comp Mole Frac (Propane) 0.015 0.015 0.000 0.000 0 0.000 0 0.000 0.000 Comp Mole Frac (i-Butane) 0.01 0.01 0.000 0.000 0 0.000 0 0.000 0.000 Comp Mole Frac (i-Pentane) 0.005 0.005 0.003 0.003 0 0.003 0 0.001 0.001 Comp Mole Frac (CO) 0 0 0.000 0.000 0 0.000 0 0.282 0.282 Comp Mole Frac (Hydrogen) 0 0 0.033 0.033 0 0.033 0 0.503 0.503 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 32 The green box indicated in table 6-1 highlights the mole fraction of CO and Hydrogen gas in the outlet stream of the ATR. The ratio of H2:CO which was determined based on this result is 1.8. It is also necessary to look into the process conditions of the inlet and outlet streams of the system to determine energy usage and the yields of the products of the simulated configuration. In table 6-2, all the process conditions of all relevant streams are specified. The process condition of all three configurations is the same. Reactor data is also shown in table 6-3. Table 6-2: Process condition data of configuration 1 Unit NG TO PREREFORMER OUT PREREFORMER STEAM IN ATR Oxygen OUT ATR Vapour Fraction 1 1 1 1 1 1 1 Temperature C 40 455 388 252 675 200 1030 Pressure kPa 3000 2900 2900 4045 2800 3000 2600 Molar Flow kgmole/h 22000 22000 34488 11000 34488 12038 80271 Mass Flow kg/h 383803 383803 581970 198166 581970 385204 967169 Liquid Volume Flow m3/h 1221 1221 1486 199 1486 339 2682 Heat Flow kJ/h -1.68E+09 -1.21E+09 -3.80E+09 -2.60E+09 -3.24E+09 6.20E+07 -3.18E+09 Table 6-3: ATR (CSTR) data ATR (CSTR) Vessel Temperature 1030 °C Vessel Pressure 2800 kPa Vapor Molar Flow 3.12E+04 kgmole/h Heat Flow -9.81E+07 kJ/h Act.% Conversion_1 100 Act.% Conversion_2 100 Act.% Conversion _3 29.52 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 39 6.2 RESULT OF FISCHER-TROPSCH REACTOR OPTIMIZATION In section 6.2, the results of all case studies pertaining to the FTR are presented. As mentioned in chapter 5, there are 2 case studies involving FTR which are case 3 and 4. Case 3 and case 4 study the effect of void fraction in FTR and the effect of staging of FTR respectively. Table 6-11 shows the product distribution with varying void fraction. Case 3: Influence of cobalt catalyst characteristic on product distribution Table 6-13: Variation of molar fraction of desired product with catalyst void fraction in FTR Name/VOID 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 Comp Mole Frac (Methane) 1.51E-2 2.64E-2 3.45E-2 4.09E-2 4.44E-2 4.75E-2 4.84E-2 5.00E-2 5.21E-2 5.52E-2 Comp Mole Frac (Ethane) 1.18E-5 1.32E-5 1.39E-5 1.47E-5 1.46E-5 1.49E-5 1.44E-5 1.45E-5 1.48E-5 1.57E-5 Comp Mole Frac (Propane) 5.98E-6 6.54E-6 8.19E-6 8.26E-6 6.95E-6 8.76E-6 7.11E-6 6.80E-6 7.65E-6 8.83E-6 Comp Mole Frac (iButane) 1.31E-5 1.42E-5 1.76E-5 1.78E-5 1.49E-5 1.88E-5 1.52E-5 1.45E-5 1.63E-5 1.89E-5 Comp Mole Frac (iPentane) 2.58E-2 2.86E-2 3.01E-2 3.16E-2 3.11E-2 3.15E-2 2.99E-2 2.99E-2 3.08E-2 3.29E-2 Comp Mole Frac (nHexane) 1.56E-4 1.66E-4 2.04E-4 2.05E-4 1.70E-4 2.14E-4 1.71E-4 1.64E-4 1.84E-4 2.14E-4 Comp Mole Frac (nHeptane) 3.87E-4 4.06E-4 4.97E-4 4.97E-4 4.13E-4 5.18E-4 4.14E04 3.96E-4 4.45E-4 5.17E-4 Comp Mole Frac (nOctane) 1.26E-3 1.31E-3 1.58E-3 1.57E-3 1.31E-3 1.63E-3 1.30E-3 1.24E-3 1.39E-3 1.61E-3 Comp Mole Frac (nNonane) 2.81E-3 2.86E-3 3.45E-3 3.42E-3 2.85E-3 3.54E-3 2.83E-3 2.70E-3 3.03E-3 3.50E-3 Comp Mole Frac (nDecane) 8.40E-3 8.57E-3 1.02E-2 1.01E-2 8.43E-3 1.04E-2 8.30E-3 7.90E-3 8.84E-3 1.02E-2 Comp Mole Frac (nC11) 1.75E-2 1.78E-2 2.07E-2 2.06E-2 1.73E-2 2.08E-2 1.67E-2 1.59E-2 1.78E-2 2.06E-2 Comp Mole Frac (nC12) 3.82E-2 3.97E-2 4.48E-2 4.50E-2 3.84E-2 4.41E-2 3.54E-2 3.36E-2 3.75E-2 4.40E-2 Comp Mole Frac (nC13) 7.61E-2 8.27E-2 8.84E-2 9.10E-2 7.96E-2 8.47E-2 6.78E-2 6.44E-2 7.15E-2 8.61E-2 Comp Mole Frac (nC14) 1.10E-1 1.25E-1 1.28E-1 1.35E-1 1.20E-1 1.19E-1 9.51E-2 9.02E-2 1.00E-1 1.24E-1 Comp Mole Frac (nC15) 8.30E-2 9.55E-2 9.49E-2 1.01E-1 9.13E-2 8.80E-2 7.00E-2 6.64E-2 7.36E-2 9.17E-2 Comp Mole Frac (nC16) 8.79E-2 1.02E-1 1.01E-1 1.08E-1 9.79E-2 9.29E-2 7.37E-2 6.99E-2 7.76E-2 9.72E-2 Comp Mole Frac (nC17) 8.93E-2 1.04E-1 1.02E-1 1.10E-1 9.99E-2 9.43E-2 7.49E-2 7.10E-2 7.88E-2 9.89E-2 Comp Mole Frac (nC18) 9.01E-2 1.06E-1 1.03E-1 1.11E-1 1.01E-1 9.51E-2 7.55E-2 7.15E-2 7.94E-2 9.98E-2 Comp Mole Frac (nC19) 9.04E-2 1.06E-1 1.04E-1 1.12E-1 1.01E-1 9.54E-2 7.57E-2 7.18E-2 7.96E-2 1.00E-1 Comp Mole Frac (nC20) 4.62E-2 5.36E-2 5.22E-2 5.62E-2 5.10E-2 4.79E-2 3.80E-2 3.61E-2 4.00E-2 5.03E-2 Comp Mole Frac (nC30) 1.90E-1 7.52E-2 6.10E-2 4.36E-3 9.68E-2 1.07E-1 2.72E-1 3.04E-1 2.35E-1 7.09E-2 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 40 The product was classified into three groups depending on the length of carbon chain namely 𝐶5 → 10, 𝐶11 → 20 and 𝐶21+. This was done to make the analysis process easier and it is also consistent with this study’s objective which is to increase molar fraction of desired product. A graph of multiple hydrocarbon products distribution variation with void fraction was plotted as shown in figure 6-1. Figure 6-1: variation of molar fraction of desired product with catalyst void fraction in FTR It was observed that among all catalyst parameters mentioned in section 5.2, only void fraction shows influence on the product distribution. Other parameters however did not give any effect. Hence, only void fraction parameter was pursued and studied. The analysis of this result will be discussed in the following part of this thesis. Table 6-14 shows the summary of the reaction data in the FTR for case 3. Table 6-14: single stage FTR data FTR Reactor Volume (m3) 2000 Bed Voidage 0.4 Act.Conversion (%) 19.65 0.0000 0.1000 0.2000 0.3000 0.4000 0.5000 0.6000 0.7000 0.8000 0.9000 1.0000 0246810 12 MOLE FRACTION VOID FRACTION Product Distribution C5-10 C11-20 C21+ Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 41 Case 4: Effect of staging of FTR with Hydrogen dossification The effect of staging of FTR on overall conversion and product distribution was investigated as shown in figure 5-7. In the second and third FTR, pure hydrogen was injected to maintain the H2:CO ratio at every inlet of the FTR. The composition of the products and the reactor data are shown in table 6-15 and table 6-16 respectively. Table 6-17 shows the operation condition and molar flows of streams. Table 6-15: Molar fraction data of multiple stage FTR simulation IN FTR 1 OUT FTR 1 IN FTR 2 OUT FTR 2 UNCONV. GAS PRODUCT 2 Comp Mole Frac (H2O) 0.000 0.034 0.033 0.061 0.002 0.001 Comp Mole Frac (Oxygen) 0.000 0.000 0.000 0.000 0.000 0.000 Comp Mole Frac (Methane) 0.003 0.037 0.037 0.064 0.068 0.008 Comp Mole Frac (Ethane) 0.000 0.000 0.000 0.000 0.000 0.000 Comp Mole Frac (Propane) 0.000 0.000 0.000 0.000 0.000 0.000 Comp Mole Frac (i-Butane) 0.000 0.000 0.000 0.000 0.000 0.000 Comp Mole Frac (i-Pentane) 0.001 0.001 0.001 0.001 0.001 0.020 Comp Mole Frac (n-Hexane) 0.000 0.000 0.000 0.000 0.000 0.000 Comp Mole Frac (n-Heptane) 0.000 0.000 0.000 0.000 0.000 0.000 Comp Mole Frac (n-Octane) 0.000 0.000 0.000 0.000 0.000 0.002 Comp Mole Frac (n-Nonane) 0.000 0.000 0.000 0.000 0.000 0.003 Comp Mole Frac (n-Decane) 0.000 0.000 0.000 0.000 0.000 0.010 Comp Mole Frac (n-C11) 0.000 0.000 0.000 0.000 0.000 0.021 Comp Mole Frac (n-C12) 0.000 0.000 0.000 0.000 0.000 0.045 Comp Mole Frac (n-C13) 0.000 0.000 0.000 0.000 0.000 0.088 Comp Mole Frac (n-C14) 0.000 0.000 0.000 0.000 0.000 0.126 Comp Mole Frac (n-C15) 0.000 0.000 0.000 0.000 0.000 0.095 Comp Mole Frac (n-C16) 0.000 0.000 0.000 0.000 0.000 0.101 Comp Mole Frac (n-C17) 0.000 0.000 0.000 0.000 0.000 0.103 Comp Mole Frac (n-C18) 0.000 0.000 0.000 0.000 0.000 0.104 Comp Mole Frac (n-C19) 0.000 0.000 0.000 0.000 0.000 0.104 Comp Mole Frac (n-C20) 0.000 0.000 0.000 0.000 0.000 0.053 Comp Mole Frac (n-C30) 0.000 0.000 0.000 0.000 0.000 0.051 Comp Mole Frac (CO) 0.293 0.279 0.273 0.261 0.278 0.015 Comp Mole Frac (Hydrogen) 0.598 0.537 0.546 0.497 0.528 0.011 Comp Mole Frac (CO2) 0.105 0.112 0.109 0.115 0.122 0.039 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 42 Table 6-16: Reactor data of multiple stage FTR FTR1 FTR2 Reactor Volume (m3) 1000 1000 Heat Flow (kJ/h) 2,36E+08 1,66E+08 Bed Voidage 0.4 0.4 Act.% Conversion 11.69 9.46 Table 6-17: Process condition of multiple stage FTr simulation Unit IN FTR 1 OUT FTR 1 IN FTR 2 OUT FTR 2 UNCONV. GAS PRODUCT 2 Vapor Fraction 1 1 1 1 1 0 Temperature C 210.0 220.0 210.0 220.0 30.0 30.0 Pressure kPa 2465.0 2465.0 2465.0 2465.0 2465.0 2465.0 Molar Flow kgmole/h 33486.7 31351.7 32001.7 30437.1 28643.5 0.3 Mass Flow kg/h 473728.5 473848.3 475158.7 475186.6 442752.1 74.1 Liquid Volume Flow m3/h 1118.7 1065.5 1084.3 1045.2 1012.7 0.1 Heat Flow kJ/h -2.29E+09 -2.51E+09 -2.52E+09 -2.67E+09 -2.42E+09 -1.51E+05 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 43 6.2.1 DISCUSSION OF FISCHER-TROPSCH REACTOR OPTIMIZATION Case 3: Influence of cobalt catalyst characteristic on product distribution In FTS, the role of the catalyst is of high importance. Carbon monoxide and hydrogen must be transported from the exterior of the catalyst to the active sites. The catalyst effectiveness as well as the selectivity of the reaction is very much dependent on the presence of both reactants in about stoichiometric amount.[3] Because of this, different catalyst parameters have been investigated to learn the effect on FT product distribution. From the simulations carried out in HYSYS, only void fraction shows an effect on the conversion and product distribution. This could be because the fact that FTS is a structure insensitive reaction by the definition of Boudart.[31], [32] As depicted in Figure 6-1, selectivity also varies with void fraction. As mentioned in the objective of this study, the aim is to maximize hydrocarbon product of range C13− C19 which can be represented by the line C11− C20 in figure 6-1. The optimal void fraction to maximize the desired product was found to be 0.4. This means that 60% of the reactor is filled with Cobalt catalyst. We can see that selectivity towards the desired product increases steadily from 0.8 until void fraction of 0.4 baring out void fraction 1.0 and 0.9 since it is impossible to actually produce FT products without catalyst or with too small amount of catalyst. After this point, it can be seen that mole fraction of desired product decreases until void fraction 0.1. A dimensionless parameter( = *  ) was defined in Iglesia et al. [31] to explain the selectivity of Cobalt catalyst. This structural term  contains structural catalyst properties that can be independently measured and it appears in the  term for all values of n. Mole fraction of desired product increased because the selectivity towards it increases. This could be due to the effects of catalyst structure and site density suggested by the parameter  which gives rise to an increase in the extent of readsorption and in C5+ selectivity as Co site density is increased by increasing Co loading.[31] As seen in figure 6-1, mole fraction of desired product decreases from void fraction 0.4 to 0.1. This may be caused by the chain termination probabilities increase as  increases above a threshold value (0.4), because diffusional restrictions inhibit chain growth steps by reducing the supply of monomer species required for chain growth.[31] Therefore, we can say that the threshold and optimal value of void fraction for the desired FT product is 0.4. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 44 Case 4: Effect of staging of FTR with Hydrogen dossification In case 4, the effect of staging of FTR was investigated. The proposed configuration was a simulation of a 2-stage FTR with extra feed to be distributed after the first reactor. Pure hydrogen gas was injected into the second reactor together with the outlet stream of the first reactor. The flowsheet simulated in HYSYS can be seen in Figure 5-7. Pure hydrogen was inserted before entering the second reactor to ensure stoichiometric amount of syngas enters the following FTR. It was also observed that the consumption of reactant was not stoichiometrically equal. The FTR consumes slightly more H2 than the stoichiometric coefficient and therefore the unreacted syngas has a higher H2:CO ratio compared to the syngas inlet of the FTR. The reason for this could be because hydrogen adsorb easier onto the Cobalt catalyst surface compared to CO. Hence, more hydrogen gas is reacted. We will now compare the performance of multiple stage FTR to a single stage FTR as shown in case 3. The total reactor volume of case 3 and 4 was kept constant and the operating condition was also the same in both cases. It was observed that the total conversion of FTS in a multi stage configuration was 21.15% compared to 19.65% in a single stage FTR. This shows a 7% improvement of conversion in the multi stage reactor. Resultant of the increased in conversion, the molar fraction of the desired product increased as well. The sum of molar fraction of the desired product as shown in the red box in table 6-15 shows a 23% improvement compared to a single stage FTR configuration. We can conclude that staging of FTR reactor not only increase total conversion but it also improves the product distribution by producing a higher value product range. Table 6-18 summarizes the comparison between the two configurations of FTR. Table 6-18: Summary of multi stage FTR data Case 3 (Single FTR) Case 4 (Multiple FTR) Total Conversion (%) 19.65 21.15 Total molar fraction of desired product 0.55 0.72 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 45 6.3 RESULT OF OVERALL PROCESS OPTIMIZATION After completing all case studies, a final proposal of GTL plant configuration was made. The final configuration was simulated and determined to be the optimal configuration. Among the key criteria of the optimal configuration is as below: I. Optimal inlet temperature of ATR to improve H2:CO ratio II. Separation product and by-product after each stage of FTR III. Single stage ATR IV. Optimum Steam:Carbon ratio A graph of variation of H2:CO ratio against inlet temperature of ATR was depicted and can be seen in figure 6-2. Figure 6-2: Variation of inlet temperature of ATR with H2:CO ratio in syngas The performance of the FT reactors was also observed. Table 6-19 shows the conversion of FTR and other reactor data. Table 6-19: Reactor data of optimal FTR configuration FTR1 FTR2 Reactor Volume (m3) 1000 1000 Bed Voidage 0.4 0.4 Heat Flow (kJ/h) 2,36E+08 1,87E+08 Act.% Conversion 11.69 10.59 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 46 Lastly, it was proposed that a single stage ATR was used instead of multiple stages of ATR based on the result of previous case studies. Table 6-19 shows the H2:CO ratio at the outlet stream of ATR. Table 6-20: Molar fraction of optimal ATR configuration IN ATR OUT ATR Comp Mole Frac (H2O) 0.228 0.000 Comp Mole Frac (Oxygen) 0.000 0.000 Comp Mole Frac (Methane) 0.710 0.004 Comp Mole Frac (Ethane) 0.000 0.000 Comp Mole Frac (Propane) 0.000 0.000 Comp Mole Frac (i-Butane) 0.000 0.000 Comp Mole Frac (i-Pentane) 0.003 0.001 Comp Mole Frac (CO) 0.000 0.180 Comp Mole Frac (Hydrogen) 0.036 0.694 Comp Mole Frac (CO2) 0.022 0.121 Figure 6-3 shows the effect of steam:carbon ration on H2:CO ratio. Figure 6-3: Effect of steam on H2:CO ratio in syngas Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 47 6.3.1 DISCUSSION OF OVERALL PROCESS OPTIMIZATION After carrying out all case studies mentioned earlier, an optimal GTL plant configuration was proposed. The final configuration of the plant took into account the optimization of the ATR, FTR, Steam:Carbon ratio in the pre-reformer as well as the inlet temperature of the ATR. The mentioned parameters play an important role in optimizing the GTL plant and must be optimized accurately. The optimal configuration and condition of the ATR was proposed after studying case study 1 and 2. It was decided that the best configuration of ATR is a single stage reactor without any extra feed. This is because it was found out that staging of ATR reactor does not improve total conversion nor lower the H2:CO ratio. Hence, it is more economical to have a single stage reactor without additional heaters or coolers. Secondly, it was observed that the steam:carbon ratio in the pre-reformer affects the H2:CO ratio in syngas produced in the ATR greatly as shown in figure 6-3. By lowering the ratio of steam:carbon, the H2:CO ratio can be improved greatly which would be useful to use in FTS. It was decided that a steam:carbon ratio of 0.4 should be used in the pre-reformer. This ratio is ideal as it would reduce H2:CO ratio in syngas and at the same time avoid soot formation. Next, as shown in figure 6-2, increasing outlet temperature in ATR could lower the H2:CO ratio in the resultant syngas. Excessive temperature however would not be possible due to material and catalyst limitation. The recommended inlet temperature of ATR was decided to be around 675 °C. This will help to produce a low H2:CO ratio of syngas. Ideally, a better kinetic model is needed to model the reaction in the ATR to accurately optimize the reactor. Finally, it was observed that a multiple stage FTR with product separation in between the reactors could improve the overall conversion even higher compared to the configuration mentioned in case 4 and would be the optimal configuration. The process flowsheet can be seen in figure 5-8 and the result can be seen in table 6-19 and 6-20. It can be seen that the overall conversion of the optimal configuration is 22.28% compared to 21.15% in case 4. Conversion improves 5% compared to case 4. The total reactor volume is the same for all the configurations but by removing these streams an increase in residence time occurs. Higher residence time will increase the conversion of reactants (CO+H2) resulting in higher production amounts. This configuration however would involve a higher capital cost as it needs additional coolers and 3-phase separators. A cost-benefit study should be carried out to determine the preferred configuration. Taking everything into account, a final and optimal GTL plant flow diagram has been proposed and can be seen in figure 6-4. Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 48 Figure 6-4: Optimal GTL plant configuration Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 55 Column1 NG TO PRE-REFORMER DUMMY-1 OUT PRE-REFORMER STEAM IN ATR Oxygen OUT ATR 8 9 Water out IN FTR OUT FTR INSEPERATOR UNCONVERTED GAS CRUDE PRODUCT WATER Comp Mole Frac (H2O) 0 0 0.33296 0.332633 1 0.332633 0 0 0 0 0 0 0.065873 0.065873 0.002206 0.000602 0.998326 Comp Mole Frac (Oxygen) 0 0 0 0 0 0 1 0.000007 0.000007 0.000007 0.000007 0.000007 0.000007 0.000007 0.000008 0.000001 0 Comp Mole Frac (Methane) 0.95 0.95 0.610514 0.610826 0 0.610826 0 0.000044 0.000044 0.000044 0.000044 0.000044 0.084907 0.084907 0.090705 0.009254 0 Comp Mole Frac (Ethane) 0.02 0.02 0.000037 0.000037 0 0.000037 0 0.000016 0.000016 0.000016 0.000016 0.000016 0.000017 0.000017 0.000018 0.000009 0 Comp Mole Frac (Propane) 0.015 0.015 0 0 0 0 0 0 0 0 0 0 0.000004 0.000004 0.000004 0.000006 0 Comp Mole Frac (i-Butane) 0.01 0.01 0 0 0 0 0 0 0 0 0 0 0.000003 0.000003 0.000003 0.000013 0 Comp Mole Frac (i-Pentane) 0.005 0.005 0.002934 0.002935 0 0.002935 0 0.001321 0.001321 0.001321 0.001321 0.001321 0.001346 0.001346 0.001438 0.019036 0 Comp Mole Frac (n-Hexane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000003 0.000003 0.000003 0.000153 0 Comp Mole Frac (n-Heptane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.00037 0 Comp Mole Frac (n-Octane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.001188 0 Comp Mole Frac (n-Decane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.007868 0 Comp Mole Frac (n-C11) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.016464 0 Comp Mole Frac (n-C12) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.037404 0 Comp Mole Frac (n-C13) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.080394 0 Comp Mole Frac (n-C14) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.125737 0 Comp Mole Frac (n-C15) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.096318 0 Comp Mole Frac (n-C16) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.103933 0 Comp Mole Frac (n-C17) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.106353 0 Comp Mole Frac (n-C18) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.10755 0 Comp Mole Frac (n-C19) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.108017 0 Comp Mole Frac (n-C20) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.054313 0 Comp Mole Frac (n-C30) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.053396 0 Comp Mole Frac (CO) 0 0 0.000124 0.000124 0 0.000124 0 0.113275 0.113275 0.113275 0.113276 0.113275 0.030561 0.030561 0.032647 0.001594 0.000001 Comp Mole Frac (Hydrogen) 0 0 0.033433 0.033446 0 0.033446 0 0.714632 0.714632 0.714632 0.714625 0.714632 0.64332 0.64332 0.68725 0.01307 0.00002 Comp Mole Frac (CO2) 0 0 0.019998 0.019999 0 0.019999 0 0.170704 0.170704 0.170704 0.170711 0.170704 0.173941 0.173941 0.185706 0.054334 0.001653 Comp Mole Frac (n-Nonane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.002623 0 Column1 Unit NG TO PRE-REFORMER DUMMY-1 OUT PRE-REFORMER STEAM IN ATR Oxygen OUT ATR 8 9 Water out IN FTR OUT FTR INSEPERATOR UNCONVERTED GAS CRUDE PRODUCT WATER Vapour Fraction 1 1 0 1 1 1 1 1 1 1 0 1 1 0.936077 1 0 0 Temperature C40 455 376.461913 376.461913 252 675 200 1030.010013 38 38 38 210 220 30 30 30 30 Pressure kPa 3000 2900 2900 2900 4045 2800 3000 2600 2565 2565 2565 2465 2265 2230 2230 2230 2230 Molar Flow kgmole/h 8195 8195 0 13960.87867 5204 13960.87867 4448.93429 31013.72061 31013.72061 31013.72061 0 31013.72061 30436.60803 30436.60803 28491.01223 0.249253 1945.34654 Mass Flow kg/h 142966.717 142966.7174 0 236717.3833 93750.58289 236717.3833 142365.897 379081.4978 379081.4978 379081.4978 0 379081.4978 379197.9797 379197.9797 344014.984 54.389125 35128.6066 Liquid Volume Flow m3/h 454.942076 454.942076 0 573.988573 93.939775 573.988573 125.137025 1049.860869 1049.860869 1049.860869 0 1049.860869 1059.574556 1059.574556 1024.274287 0.070066 35.230203 Heat Flow kJ/h -627173957 -449480204.5 0 -1677180833 -1227700592 -1449025500 22929335.8 -1426131875 -2470180947 -2470180947 0 -2302443179 -2681842201 -2956436325 -2400004656 -111642.2 -556320027 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 56 Annex C: Case study 1 – Configuration 3 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 57 Column1 NG TO PRE-REFORMER DUMMY-1 OUT PRE-REFORMER STEAM IN ATR Oxygen OUT ATR 8 9 Water out IN FTR OUT FTR INSEPERATOR UNCONVERTED GAS CRUDE PRODUCT WATER Comp Mole Frac (H2O) 0 0 0.33296 0.332633 1 0.332633 0 0 0 0 0 0 0.065874 0.065874 0.002206 0.000603 0.998326 Comp Mole Frac (Oxygen) 0 0 0 0 0 0 1 0.000008 0.000008 0.000008 0.000008 0.000008 0.000008 0.000008 0.000008 0.000001 0 Comp Mole Frac (Methane) 0.95 0.95 0.610514 0.610826 0 0.610826 0 0.000041 0.000041 0.000041 0.000041 0.000041 0.084905 0.084905 0.090703 0.009364 0 Comp Mole Frac (Ethane) 0.02 0.02 0.000037 0.000037 0 0.000037 0 0.000016 0.000016 0.000016 0.000016 0.000016 0.000017 0.000017 0.000018 0.000009 0 Comp Mole Frac (Propane) 0.015 0.015 0 0 0 0 0 0 0 0 0 0 0.000004 0.000004 0.000004 0.000006 0 Comp Mole Frac (i-Butane) 0.01 0.01 0 0 0 0 0 0 0 0 0 0 0.000003 0.000003 0.000003 0.000013 0 Comp Mole Frac (i-Pentane) 0.005 0.005 0.002934 0.002935 0 0.002935 0 0.001321 0.001321 0.001321 0.001321 0.001321 0.001346 0.001346 0.001438 0.019409 0 Comp Mole Frac (n-Hexane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000003 0.000003 0.000003 0.000156 0 Comp Mole Frac (n-Heptane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.000378 0 Comp Mole Frac (n-Octane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.001213 0 Comp Mole Frac (n-Decane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.008028 0 Comp Mole Frac (n-C11) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.01682 0 Comp Mole Frac (n-C12) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.038382 0 Comp Mole Frac (n-C13) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.083401 0 Comp Mole Frac (n-C14) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.132223 0 Comp Mole Frac (n-C15) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.101877 0 Comp Mole Frac (n-C16) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.110349 0 Comp Mole Frac (n-C17) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.11306 0 Comp Mole Frac (n-C18) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.114408 0 Comp Mole Frac (n-C19) 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.114936 0 Comp Mole Frac (n-C20) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.057803 0 Comp Mole Frac (n-C30) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.005756 0 Comp Mole Frac (CO) 0 0 0.000124 0.000124 0 0.000124 0 0.113276 0.113276 0.113276 0.113277 0.113276 0.030561 0.030561 0.032648 0.001598 0.000001 Comp Mole Frac (Hydrogen) 0 0 0.033433 0.033446 0 0.033446 0 0.714634 0.714634 0.714634 0.714626 0.714634 0.643321 0.643321 0.687251 0.013111 0.00002 Comp Mole Frac (CO2) 0 0 0.019998 0.019999 0 0.019999 0 0.170703 0.170703 0.170703 0.17071 0.170703 0.17394 0.17394 0.185705 0.054417 0.001653 Comp Mole Frac (n-Nonane) 0 0 0 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.002678 0 Column1 Unit NG TO PRE-REFORMER DUMMY-1 OUT PRE-REFORMER STEAM IN ATR Oxygen OUT ATR 8 9 Water out IN FTR OUT FTR INSEPERATOR UNCONVERTED GAS CRUDE PRODUCT WATER Vapour Fraction 1 1 0 1 1 1 1 1 1 1 0 1 1 0.936077 1 0 0 Temperature C40 455 376.461913 376.461913 252 675 200 1030.004485 38 38 38 210 220 30 30 30 30 Pressure kPa 3000 2900 2900 2900 4045 2800 3000 2600 2565 2565 2565 2465 2265 2230 2230 2230 2230 Molar Flow kgmole/h 8195 8195 0 13960.87867 5204 13960.87867 4448.98368 31013.94173 31013.94173 31013.94173 0 31013.94173 30436.80946 30436.80946 28491.19314 0.234879 1945.38144 Mass Flow kg/h 142966.717 142966.7174 0 236717.3833 93750.58289 236717.3833 142367.478 379083.7276 379083.7276 379083.7276 0 379083.7276 379195.4806 379195.4806 344017.1469 49.09714 35129.2366 Liquid Volume Flow m3/h 454.942076 454.942076 0 573.988573 93.939775 573.988573 125.138414 1049.86626 1049.86626 1049.86626 0 1049.86626 1059.574344 1059.574344 1024.279993 0.063515 35.230835 Heat Flow kJ/h -627173957 -449480204.5 0 -1677180833 -1227700592 -1449025500 22929590.4 -1426142399 -2470189906 -2470189906 0 -2302451150 -2681851680 -2956444682 -2400013396 -101278.9371 -556330007 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 58 Annex D: Case study 2 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 59 Column1 NG TO PRE-REFORMER DUMMY-1 STEAM OUT ATR 8 9 Water out IN FTR OUT FTR INSEPERATOR UNCONVERTED GAS CRUDE PRODUCT WATER OUT ATR 2 521 IN ATR 2 OUT PRE-REFORMER IN ATR Oxygen Oxygen-2 Oxygen-2-2 IN ATR 3 OUT ATR 3 5521 Comp Mole Frac (H2O) 0 0 0.276241 1 0 0 0 0 0 0.040663 0.040663 0.00239 0.000599 0.998774 0 0 0 0.27599 0.27599 0 0 0 0 0 0 Comp Mole Frac (Oxygen) 0 0 0 0 0.000009 0.000009 0.018075 0.018075 0.018075 0.018287 0.018287 0.019017 0.001339 0.000001 0.012034 0.012034 0.012314 0 0 1 1 1 0.018075 0.018075 0.018075 Comp Mole Frac (Methane) 0.95 0.95 0.663826 0 0.000142 0.000142 0 0 0 0.052384 0.052384 0.054477 0.005151 0 0 0 0.00014 0.664068 0.664068 0 0 0 0 0 0 Comp Mole Frac (Ethane) 0.02 0.02 0.00005 0 0.000022 0.000022 0.000021 0.000021 0.000021 0.000021 0.000021 0.000022 0.00001 0 0.000021 0.000021 0.000021 0.00005 0.00005 0 0 0 0.000021 0.000021 0.000021 Comp Mole Frac (Propane) 0.015 0.015 0 0 0 0 0 0 0 0.000004 0.000004 0.000004 0.000006 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (i-Butane) 0.01 0.01 0 0 0 0 0 0 0 0.000003 0.000003 0.000003 0.000012 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (i-Pentane) 0.005 0.005 0.003189 0 0.00137 0.00137 0.001345 0.001345 0.001345 0.001361 0.001361 0.001415 0.017896 0 0.001353 0.001353 0.001353 0.00319 0.00319 0 0 0 0.001345 0.001345 0.001345 Comp Mole Frac (n-Hexane) 0 0 0 0 0 0 0 0 0 0.000003 0.000003 0.000003 0.000148 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-Heptane) 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.000362 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-Octane) 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.001168 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-Decane) 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.007819 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C11) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.016507 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C12) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.038019 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C13) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0.083849 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C14) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.134414 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C15) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.103979 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C16) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.112831 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C17) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.115651 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C18) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.117046 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C19) 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.117588 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C20) 0 0 0 0 0 0 0 0 0 0 0 0 0.059154 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (n-C30) 0 0 0 0 0 0 0 0 0 0 0 0 0.007836 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (CO) 0 0 0.000189 0 0.151942 0.151942 0.148876 0.148877 0.148876 0.098232 0.098232 0.102157 0.00457 0.000003 0.149792 0.149792 0.150073 0.000189 0.000189 0 0 0 0.148876 0.148876 0.148876 Comp Mole Frac (Hydrogen) 0 0 0.035119 0 0.704015 0.704015 0.691379 0.691376 0.691379 0.647079 0.647079 0.672932 0.011678 0.000017 0.695633 0.695633 0.695353 0.035128 0.035128 0 0 0 0.691379 0.691379 0.691379 Comp Mole Frac (CO2) 0 0 0.021387 0 0.142499 0.142499 0.140303 0.140306 0.140303 0.141944 0.141944 0.147567 0.039782 0.001204 0.141166 0.141166 0.140746 0.021386 0.021386 0 0 0 0.140303 0.140303 0.140303 Comp Mole Frac (n-Nonane) 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.002586 0 0 0 0 0 0 0 0 0 0 0 0 Column1 Unit NG TO PRE-REFORMER DUMMY-1 STEAM OUT ATR 8 9 Water out IN FTR OUT FTR INSEPERATOR UNCONVERTED GAS CRUDE PRODUCT WATER OUT ATR 2 521 IN ATR 2 OUT PRE-REFORMER IN ATR Oxygen Oxygen-2 Oxygen-2-2 IN ATR 3 OUT ATR 3 5521 Vapour Fraction 1 1 0 1 1 1 1 0 1 1 0.961581 1 0 0 1 1 1 1 1 1 1 1 1 1 1 Temperature C40 455 387.533737 252 1029.983868 700 38 38 210 220 30 30 30 30 697.605655 700 694.067106 387.533737 675 200 200 200 697.052805 697.054321 38 Pressure kPa 3000 2900 2900 4045 2600 2565 2365 2365 2265 2065 2030 2030 2030 2030 2465 2465 2565 2900 2800 3000 3000 3000 2465 2365 2365 Molar Flow kgmole/h 22000 22000 011000 80271.08476 80271.08476 81771.22174 0 81771.22174 80826.09573 80826.09573 77720.82529 0.628881 3104.64157 81271.22174 81271.22174 81271.08476 34488.1329 34488.1329 12037.617 1000 500 81771.22174 81771.22174 81771.22174 Mass Flow kg/h 383803.2683 383803.2683 0 198166.1053 967169.3451 967169.3451 1015171.209 01015171.209 1015477.738 1015477.738 959316.933 133.931317 56026.874 999171.2093 999171.2093 999169.3451 581969.5863 581969.5863 385203.743 32000 16000 1015171.209 1015171.209 1015171.209 Liquid Volume Flow m3/h 1221.32101 1221.32101 0 198.566013 2681.655244 2681.655244 2724.281164 0 2724.281164 2740.366435 2740.366435 2684.017422 0.173086 56.175927 2710.217456 2710.217456 2709.78266 1486.434875 1486.434875 338.587058 28.127416 14.063708 2724.281164 2724.281164 2724.281164 Heat Flow kJ/h -1683688476 -1206658267 0 -2595062742 -3182056774 -4117420714 -5855352888 0 -5418290073 -6030609265 -6656185063 -5768208811 -273017.1674 -887703234 -4112399735 -4105691523 -4112266820 -3801720827 -3244003819 62040602.1 5153894.02 2576947.01 -4103114576 -4103214685 -5855352888 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 60 Annex E: Case study 3 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 61 Column1 Unit NG TO PRE-REFORMER DUMMY-1 STEAM OUT ATR IN FTR OUT FTR 1 OUT PRE-REFORMER IN ATR Oxygen IN FTR 1 OUT ATR HEATER IN HEATER IN SYSTEM CO INLET ATR TO SEPARATOR UNREACTED GAS CRUDE PROD WATER Vapour Fraction 1 1 0 1 1 1 1 1 1 1 1 1 1 1 1 1 0.936718 1 0 0 Temperature C40 455 397.161071 252 1029.999606 210 220 397.161071 675 200 209.983729 38 38 40 210 596.028745 30 30 30 30 Pressure kPa 3000 2900 2900 4045 2665 2665 2665 2900 2800 3000 2665 2665 2665 3000 2665 2800 2665 2665 2665 2665 Molar Flow kgmole/h 8195 8195 03278 28886.64785 28886.64785 29634.32088 12018.47155 12018.47155 4464.35817 33486.64785 28886.64785 28886.64785 8195 4600 16482.82971 29634.32088 27758.98795 0.342099 1874.99083 Mass Flow kg/h 142966.717 142966.7174 0 59053.4994 344878.2339 344878.2339 473870.403 202020.2953 202020.2953 142859.461 473728.3762 344878.2339 344878.2339 142966.7174 128850.142 344879.7567 473870.403 439938.4237 90.398303 33841.581 Liquid Volume Flow m3/h 454.942076 454.942076 0 59.172672 957.536722 957.536722 1022.627366 538.530366 538.530366 125.570859 1118.722707 957.536722 957.536722 454.942076 161.185985 664.101225 1022.627366 988.579348 0.114484 33.933534 Heat Flow kJ/h -627173957 -449480204.5 0 -773328697 -1007543592 -1805446156 -2693027947 -1222808899 -1030487273 23008828.9 -2289393412 -1958979256 -1958979256 -627173957.2 -483947256 -1007478444 -2956027070 -2419735198 -180936.0013 -536110936 NG TO PRE-REFORMER DUMMY-1 STEAM OUT ATR IN FTR OUT FTR 1 OUT PRE-REFORMER IN ATR Oxygen IN FTR 1 OUT ATR HEATER IN HEATER IN SYSTEM CO INLET ATR TO SEPARATOR UNREACTED GAS CRUDE PROD WATER Comp Mole Frac (H2O) 0 0 0.227817 1 0 0 0.065012 0.227632 0.227632 0 0 0 0 0 0 0.165978 0.065012 0.001949 0.000594 0.998663 Comp Mole Frac (Oxygen) 0 0 0 0 0 0 0 0 0 1 0 0 0 0 0 0.270849 0 0 0 0 Comp Mole Frac (Methane) 0.95 0.95 0.71 0 0.003503 0.003503 0.068427 0.710181 0.710181 0 0.003022 0.003503 0.003503 0.95 0 0.517829 0.068427 0.07305 0.008337 0 Comp Mole Frac (Ethane) 0.02 0.02 0.000066 0 0.000028 0.000028 0.000027 0.000066 0.000066 0 0.000024 0.000028 0.000028 0.02 0 0.000048 0.000027 0.000029 0.000015 0 Comp Mole Frac (Propane) 0.015 0.015 0 0 0 0 0.000003 0 0 0 0 0 0 0.015 0 0 0.000003 0.000003 0.000006 0 Comp Mole Frac (i-Butane) 0.01 0.01 0 0 0 0 0.000003 0 0 0 0 0 0 0.01 0 0 0.000003 0.000003 0.000012 0 Comp Mole Frac (i-Pentane) 0.005 0.005 0.003409 0 0.001418 0.001418 0.001383 0.003409 0.003409 0 0.001224 0.001418 0.001418 0.005 0 0.002486 0.001383 0.001476 0.020027 0 Comp Mole Frac (n-Hexane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000003 0.000138 0 Comp Mole Frac (n-Heptane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000337 0 Comp Mole Frac (n-Octane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.001089 0 Comp Mole Frac (n-Decane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.007041 0 Comp Mole Frac (n-C11) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.014434 0 Comp Mole Frac (n-C12) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.03122 0 Comp Mole Frac (n-C13) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.061876 0 Comp Mole Frac (n-C14) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0 0.089432 0 Comp Mole Frac (n-C15) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0 0.06772 0 Comp Mole Frac (n-C16) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0 0.071813 0 Comp Mole Frac (n-C17) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0 0.073087 0 Comp Mole Frac (n-C18) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0 0.073715 0 Comp Mole Frac (n-C19) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0 0.073959 0 Comp Mole Frac (n-C20) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.037857 0 Comp Mole Frac (n-C30) 0 0 0 0 0 0 0.000003 0 0 0 0 0 0 0 0 0 0.000003 0 0.294634 0 Comp Mole Frac (CO) 0 0 0.000271 0 0.180254 0.180254 0.26592 0.000271 0.000271 0 0.292861 0.180254 0.180254 0 1 0.000197 0.26592 0.283884 0.016228 0.000012 Comp Mole Frac (Hydrogen) 0 0 0.036012 0 0.693637 0.693637 0.4811 0.036018 0.036018 0 0.598353 0.693637 0.693637 0 0 0.026263 0.4811 0.513601 0.011491 0.000018 Comp Mole Frac (CO2) 0 0 0.022424 0 0.121159 0.121159 0.118102 0.022422 0.022422 0 0.104516 0.121159 0.121159 0 0 0.016349 0.118102 0.125992 0.042562 0.001307 Comp Mole Frac (n-Nonane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.002375 0 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 62 Annex F: Case study 4 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 63 Column1 NG TO PRE-REFORMER DUMMY-1 STEAM OUT ATR IN FTR OUT FTR 1 OUT PRE-REFORMER IN ATR Oxygen IN FTR 1 OUT ATR HEATER IN HEATER IN SYSTEM CO H2 IN FTR 3 OUT FTR 2 IN SEP 2 V2 PRODUCT 2 WATER 2 V1 HOT Comp Mole Frac (H2O) 0 0 0.227817 1 0 0 0.036756 0.227632 0.227632 0 0 0 0 0 0 0 0.036005 0.06508 0.06508 0.00207 0.000599 0.998777 0.036756 Comp Mole Frac (Oxygen) 0 0 0 0 0 0 0 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 Comp Mole Frac (Methane) 0.95 0.95 0.71 0 0.003507 0.003507 0.040003 0.710181 0.710181 0 0.003025 0.003507 0.003507 0.95 0 0 0.039187 0.068434 0.068434 0.073053 0.008204 0 0.040003 Comp Mole Frac (Ethane) 0.02 0.02 0.000066 0 0.000028 0.000028 0.000026 0.000066 0.000066 0 0.000024 0.000028 0.000028 0.02 0 0 0.000025 0.000026 0.000026 0.000028 0.000015 0 0.000026 Comp Mole Frac (Propane) 0.015 0.015 0 0 0 0 0.000003 0 0 0 0 0 0 0.015 0 0 0.000003 0.000004 0.000004 0.000004 0.000008 0 0.000003 Comp Mole Frac (i-Butane) 0.01 0.01 0 0 0 0 0.000003 0 0 0 0 0 0 0.01 0 0 0.000003 0.000004 0.000004 0.000004 0.000017 0 0.000003 Comp Mole Frac (i-Pentane) 0.005 0.005 0.003409 0 0.001418 0.001418 0.001314 0.003409 0.003409 0 0.001224 0.001418 0.001418 0.005 0 0 0.001287 0.001357 0.001357 0.001448 0.020568 0 0.001314 Comp Mole Frac (n-Hexane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000003 0.000003 0.000003 0.000198 0 0.000002 Comp Mole Frac (n-Heptane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000003 0.000003 0.000003 0.000487 0 0.000002 Comp Mole Frac (n-Octane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000003 0.000003 0.000003 0.001567 0 0.000002 Comp Mole Frac (n-Decane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.000002 0.010162 0 0.000002 Comp Mole Frac (n-C11) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000002 0.000002 0.000002 0.02084 0 0.000001 Comp Mole Frac (n-C12) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000002 0.000002 0.000001 0.044996 0 0.000001 Comp Mole Frac (n-C13) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000002 0.000002 0.000001 0.088996 0 0.000001 Comp Mole Frac (n-C14) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000002 0.000002 0 0.128806 0 0.000001 Comp Mole Frac (n-C15) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0 0.097649 0 0.000001 Comp Mole Frac (n-C16) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0 0.103609 0 0.000001 Comp Mole Frac (n-C17) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0 0.105469 0 0.000001 Comp Mole Frac (n-C18) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0 0.106395 0 0.000001 Comp Mole Frac (n-C19) 0 0 0 0 0 0 0.000001 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0.000001 0 0.106759 0 0.000001 Comp Mole Frac (n-C20) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.000001 0.000001 0 0.054692 0 0 Comp Mole Frac (n-C30) 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.031345 0 0 Comp Mole Frac (CO) 0 0 0.000271 0 0.18025 0.18025 0.277623 0.000271 0.000271 0 0.292858 0.18025 0.18025 0 1 0 0.271956 0.259585 0.259585 0.277105 0.014861 0.000011 0.277623 Comp Mole Frac (Hydrogen) 0 0 0.036012 0 0.693634 0.693634 0.532053 0.036018 0.036018 0 0.598351 0.693634 0.693634 0 0 1 0.541605 0.489615 0.489615 0.522661 0.010959 0.000016 0.532053 Comp Mole Frac (CO2) 0 0 0.022424 0 0.121162 0.121162 0.112199 0.022422 0.022422 0 0.104519 0.121162 0.121162 0 0 0 0.109909 0.11587 0.11587 0.12361 0.039366 0.001196 0.112199 Comp Mole Frac (n-Nonane) 0 0 0 0 0 0 0.000002 0 0 0 0 0 0 0 0 0 0.000002 0.000002 0.000002 0.000002 0.003435 0 0.000002 Column1 Unit NG TO PRE-REFORMER DUMMY-1 STEAM OUT ATR IN FTR OUT FTR 1 OUT PRE-REFORMER IN ATR Oxygen IN FTR 1 OUT ATR HEATER IN HEATER IN SYSTEM CO H2 IN FTR 3 OUT FTR 2 IN SEP 2 V2 PRODUCT 2 WATER 2 V1 HOT Vapour Fraction 1 1 0 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 0.936771 1 0 0 1 Temperature C40 455 397.161071 252 1030.018216 210 220 397.161071 675 200 209.983731 38 38 40 210 210 209.97923 220 30 30 30 30 210 Pressure kPa 3000 2900 2900 4045 2665 2665 2665 2900 2800 3000 2665 2665 2665 3000 2665 2665 2665 2665 2465 2465 2465 2465 2665 Molar Flow kgmole/h 8195 8195 03278 28886.42378 28886.42378 31194.13721 12018.47155 12018.47155 4464.35817 33486.42378 28886.42378 28886.42378 8195 4600 650 31844.13721 30205.87912 30205.87912 28295.9878 0.328221 1909.5631 31194.13721 Mass Flow kg/h 142966.717 142966.7174 0 59053.4994 344877.7577 344877.7577 473843.4261 202020.2953 202020.2953 142859.461 473727.9 344877.7577 344877.7577 142966.7174 128850.142 1310.40002 475153.8262 475180.8594 475180.8594 440650.974 69.852516 34460.0329 473843.4261 Liquid Volume Flow m3/h 454.942076 454.942076 0 59.172672 957.533421 957.533421 1061.576872 538.530366 538.530366 125.570859 1118.719406 957.533421 957.533421 454.942076 161.185985 18.757757 1080.334628 1039.429092 1039.429092 1004.787237 0.090224 34.551631 1061.576872 Heat Flow kJ/h -627173957 -449480204.5 0 -773328697 -1007543827 -1805463998 -2525398177 -1222808899 -1030487273 23008828.9 -2289411254 -1958996347 -1958996347 -627173957.2 -483947256 3457086.97 -2532010005 -2697806943 -2965288806 -2419166556 -142522.178 -545979727 -2535467092 Optimization Study of Fishcer-Tropsch Gas-to-Liquid (GTL) process using Aspen HYSYS 64 Annex G: Final Configuration