Fischer Tropsch Catalyst Structures & Process Design for JP-5 Fuel Integrated with MFEC by Tunde B.A Dokun A thesis submitted to the Graduate Faculty of Auburn University in partial fulfillment of the requirements for the Degree of Master of Science Auburn, Alabama August 9, 2010 Keywords: Process and Product Design, Gas-to-Liquids Fischer Tropsch Synthesis, Packed Bed, Aspen Simulation, MFEC Copyright 2010 by Tunde Dokun Approved by Bruce J. Tatarchuck, Chair, Professor of Chemical Engineering Mario R. Eden, Associate Professor of Chemical Engineering Robert W. Ashurst, Assistant Professor of Chemical Engineering ii Abstract Modern societies are seeking new sources and corresponding technology as traditional fossil fuels are becoming more difficult to access because of their remote locations. A potentially attractive solution is to convert natural gas to a synthesis gas and then synthesize longer-chain hydrocarbons through a Fischer-Tropsch (FTS) reaction. However, conventional FTS reactors are faced with the challenge of heat removal due to the reaction being highly exothermic. A design framework has been established, which incorporates functionality of the FTS catalyst and its impact on the FTS balance of plant (BOP). Process simulation tools are used in an iterative design to develop a unique plant optimized for the production of JP-5 and other value-added hydrocarbons under process constraints and size limitations. The focus of this work is to evaluate the effects of utilizing a micro-fibrous entrapped catalyst (MFEC) on FTS plant scalability, physical plant design, and critical front-end capital cost in order to design a mobilized Fischer-Tropsch process for the purpose of producing JP- 5. The product distribution of the FTS reaction is given by the Anderson-Schulz-Flory distribution. MFEC is comprised of a small grain catalyst entrapped within a sinter-locked network of a metal. With the use of this metal fiber network, we are able to increase the effective thermal conductivity within the reactor by about 90%, and reduce the temperature rise within the reactor as compared to a conventional Packed bed reactor. MFEC are readily manufactured and provide high intra-particle and mass transport properties. The temperature uniformity within the reactor will ensure and enhance JP-5 selectivity. iii Using an implicit finite integrating scheme, we are able to determine the reactor temperature profile by modeling a plug flow reactor which includes the energy balances on the gas phase. The catalyst temperature will be assumed to be uniform inside the catalyst particles, so all heat of reaction is generated inside the catalyst particles. Moreover, with Aspen TM integrated simulation, we are able to simulate an overall GTL plant capable of producing 500 bpd of JP-5. The simulation takes into account a vapor-liquid relationship of the different FTS hydrocarbon products with the intrinsic kinetics of the FTS and a water gas shift reaction for a promoted Iron catalyst, Fe/Cu/K as a basis of development. The application of applied mathematics and numerical methods in solving the sets of differential equations is the key to fully understanding the temperature profile. By effectively designing an FTS reactor that has an effective heat transfer mechanism due to a high effective intra-bed thermal conductivity, a better control on intra-bed hot spots and product selectivity is achieved. A High selectivity towards JP-5 will be observed and thus, we are able to reduce the balance of plant (BOP) requirements for the FTS reactor and downstream operations. iv Acknowledgements The project undertaken for this research thesis has involved collaboration not only between academia and industry, but among a vast number of educational disciplines, family and friends. Due to the broad interdisciplinary nature of this project, there are many people in many walks of life that have contributed to the success of the project, and I would like to recognize and thank all of them. First, I would first like to thank my parents; Tele and Tope Dokun for their unconditional love and encouraging support throughout my life. I would like to thank my major professor, Dr. Bruce Tatarchuk, research committee; Dr. Mario Eden, and Dr. Bob Ashurst for their guidance and constructive feedback during the completion of this project. I would like to thank Ms. Karen Cochran and Ms. Kim Dennis for their support and services with academic and non-academic related issues. I would also like to thank the expertise of co-authors during my publications and presentations in relation to the project, particularly Dr. Donald Cahela, Dr. Hongyun Yang and Dr. Norm Sammons. Finally, I would like to thank all of my office mates and collaborators; Min Sheng, Josh Jackson, Pavielle Lockhart and Steve Saunders for their ability to provide alternative viewpoints needed to tackle the problem from all angles. To each of you ? My sincerest gratitude and love. v Table of Contents Abstract ............................................................................................................................ ii Acknowledgements .......................................................................................................... iv List of Tables ................................................................................................................... vii List of Figures .................................................................................................................. ix List of Abbreviations ....................................................................................................... xi Chapter 1 ? Introduction .................................................................................................. 1 Chapter 2 ? Theoretical Background ............................................................................... 3 2.1 History ..................................................................................................... 3 2.2 FTS Catalyst............................................................................................. 3 2.2.1 Product Selectivity .......................................................................... 6 2.3 FTS Reactor ............................................................................................. 8 2.3.1 Reactor Design ............................................................................. 9 2.3.2 Fixed Bed ..................................................................................... 11 2.3.3 Fluidized Bed ............................................................................... 12 2.3.4 Slurry Bed .................................................................................... 12 2.4 FTS Polymerization Reaction .................................................................. 14 2.5 FTS Mechanism/Kinetics......................................................................... 14 2.5.1 Fischer - Tropsch Synthesis ........................................................ 14 2.5.2 Water- Gas Shift Reaction .......................................................... 18 vi Chapter 3 ? FTS Catalyst Design for Jet Fuel ................................................................. 20 3.1 Catalyst Design and Optimization ........................................................... 21 3.2 Catalyst Matrix......................................................................................... 24 3.2.1 Catalyst for Jet Fuel Production .................................................... 26 3.2.2 Phase Changes during Synthesis ................................................. 35 3.2.3 Influence of Promoters on Physical Properties and Synthesis ..... 36 3.2.4 Catalyst Size and Shape .............................................................. 38 3.3 Thermal Conductivity .............................................................................. 40 Chapter 4 FTS Process Design for Jet Fuel .............................................................. 44 4.1 Introduction ............................................................................................... 45 4.2 Integrated Simulated Design ..................................................................... 51 4.2.1 Upstream ...................................................................................... 52 4.2.2 FTS Reactor ................................................................................. 57 4.2.3 Downstream ................................................................................ 65 4.3 Weight/Volume BOP Analysis ................................................................ 66 Chapter 5 ? Accomplishments and Future Directions ..................................................... 79 5.1 Accomplishments ..................................................................................... 79 5.2 Future Direction ....................................................................................... 80 Bibliography?????????????????????????????. 81 Appendices ?????????????????????????????. 88 vii List of Tables Table 2.1 Existing plants and plants under construction for Fischer-Tropsch synthesis .............. 4 Table 2.2 Summary of Catalyst, Alpha values and Operating Conditions ................................... 7 Table 2.3 Advantages and disadvantages of established reactors for FTS ................................. 13 Table 3.1 Summary of Premium Products, Catalyst and technology used ................................. 24 Table 3.2 Chain Growth Probability factors for different Catalyst ............................................ 27 Table 3.3 Estimated constants A and B ...................................................................................... 29 Table 3.4 Summary of FT Kinetics and Parameter ..................................................................... 31 Table 3.5 Characterization data of Catalyst ................................................................................ 33 Table 3.6 Optimization of Fe Catalyst Design............................................................................ 34 Table 3.7 Influence of alkali content on synthesis performance of precipitated iron catalyst .... 38 Table 3.8 Some common catalyst shape ..................................................................................... 39 Table 3.9 Thermal Conductivity and wall heat transfer coefficients .......................................... 42 Table 4.1 Selected Fuel Properties for Jet Fuel and Diesel ........................................................ 49 Table 4.2 Comparison of Syngas generation technologies (Natural Gas Feed) I ....................... 52 Table 4.3 Comparison of Syngas generation technologies. II .................................................... 53 Table 4.4 Sensitivity Analysis of steam to methane ratio and N.G temperature on GTL Performance. ............................................................................................................................... 56 Table 4.5 MFEC Reactors versus Fixed and Slurry Reactors. ................................................... 58 Table 4.6 Conventional Packed Bed Reactors vs. MFEC. ......................................................... 62 viii Table 4.7 Overall Product Yield of JP-5 . ................................................................................... 64 Table 4.8 Case A: Overall Product Yield of JP-5 - HTFT.......................................................... 69 Table 4.9 Case A: Overall Product Yield of JP-5 - LTFT. ......................................................... 70 Table 4.10 Products for Case Scenarios ..................................................................................... 74 Table 4.11 Weight and Volume of Major Process Equipments for Case A ............................... 74 Table 4.12 Weight and Volume of Major Process Equipments for Case B & C ........................ 75 Table 4.13 General Comparison Cases for 500 bpd of JP-5 ....................................................... 76 Table 4.14 Composition of Recycled Streams ............................................................................ 77 ix List of Figures Figure 2.1 Hydrocarbon Selectivity .............................................................................................. 6 Figure 2.2 Types of Fischer-Tropsch Synthesis Reactors ............................................................ 9 Figure 2.3 Selectivity of FTS for LTFT and HTFT on Fe Catalyst ............................................ 10 Figure 2.4 kinetic scheme of the successive hydrogenation of surface carbon .......................... 17 Figure 3.1 An Integrated Approach in FTS Catalyst Design ...................................................... 22 Figure 3.2 ASF Weight distributions as a function of chain growth probability ........................ 23 Figure 3.3 Influence of SiO 2 on pore volume of Fe 2 O 3 .............................................................. 25 Figure 3.4 Influence of SiO 2 on Surface Area of Fe 2 O 3 ............................................................. 25 Figure 3.5 Alpha versus Temperature for Promoted Iron ........................................................... 28 Figure 3.6 Alpha versus Syngas ratio for Promoted Iron ........................................................... 29 Figure 3.7 CO hydrogenation activities I .................................................................................... 32 Figure 3.8 CO hydrogenation activities II .................................................................................. 32 Figure 3.9 CO hydrogenation activities III ................................................................................. 33 Figure 3.10 Iron Catalyst Phase Change during FT .................................................................... 35 Figure 3.11 Influence on the unit cell size of magnetite. ............................................................ 37 Figure 3.12 The surface area of fully reduced catalyst .............................................................. 37 Figure 3.13 Effect of particle size of the catalyst on CO Conversion ........................................ 38 Figure 3.14 Catalyst Effectiveness Factor ................................................................................. 39 Figure 3.15 Diagram of Location of Thermocouples ................................................................ 41 x Figure 3.16 Temperature Bed comparisons between Packed Bed and MFEC ........................... 42 Figure 3.17 Steady-State data and segregated 2-D simulation result for heat-up of a tube packed with MFEC ..................................................................................................................... 42 Figure 4.1 U.S Energy Consumption .......................................................................................... 45 Figure 4.2 Comparison of FT Products with conventional Refinery barrel ............................... 47 Figure 4.3 Block Flow Diagram of an FT Process ..................................................................... 51 Figure 4.4 Process Flow Sheet of FTS - Upstream ..................................................................... 54 Figure 4.5 CO Conversion in FTS reactor .................................................................................. 60 Figure 4.6 Temperature profile in FTS reactor ........................................................................... 61 Figure 4.7 CO Conversion in FTS reactor .................................................................................. 61 Figure 4.8 Product Selectivity Comparisons between Packed bed and MFEC .......................... 62 Figure 4.9 Product Yield from Hydro-cracking .......................................................................... 64 Figure 4.10 Process Flow Sheet of FTS ? Downstream ............................................................. 65 Figure 4.11 Fischer Tropsch Possibilities for Mobile Skid Unit ................................................ 67 Figure 4.12 Process Flow Diagram Case A (Single Pass) .......................................................... 68 Figure 4.13 HTFT for Case A (Single Pass) ............................................................................... 69 Figure 4.14 Case A (Single Pass) Simulation validation ............................................................ 70 Figure 4.15 LTFT for Case A (Single Pass) ............................................................................... 71 Figure 4.16 Process Flow Diagram for Case B (Single Pass with Hydro-cracking unit) ........... 72 Figure 4.17 Process Flow Diagram for Case C (Single Pass with Hydro-cracking unit/Lt. Naphta) ........................................................................................................................................ 72 Figure 4.18 Reactor Productivity ................................................................................................ 77 xi Abbreviations ? Biot number, rtw kRh / p C fluid specific heat G superficial mass flowrate Nu Apparent wall Nusselt number p d Particle Diameter fw Nu Apparent fluid/wall Nusselt number fpwf kdh / RF Pe Radial fluid Peclet number, rfpp kdGC / Re Reynolds number, ?/ p Gd Pr Pradtl number, gp kC /? rg k radial bulk thermal conductivity of gas rs k radial bulk thermal conductivity of solid rf k fluid bulk thermal conductivity wf h wall/fluid heart transfer coefficient fs h fluid/solid heat transfer coefficient ws h wall/solid heat transfer coefficient Bi Biot number gw krh / U Overall heat transfer coefficient w h wall heat transfer coefficient r k Effective thermal conductivity s u superficial mass velocity A r rate of reaction of CO ? voidage ? tap tapped density, gm/cc ? bulk bulk density, gm/cc ? part particle density, gm/cc L Catalyst loading, gm cat/total gm V p pore volume of support material, cc/gm 1 ?Research is to see what everybody else sees, and to think what nobody else has thought.? Albert Szent-Gyorgyi Chapter 1 Introduction 1.1 Overview The United States Department of Defense (DoD) is interested in low-sulfur, environmentally clean Fischer-Tropsch fuels?specifically Jet fuel, JP-5,?due to the inherent security risks associated with and the instability in the crude oil supply. Currently, the world?s fuel and chemical production is based predominantly on petroleum crude oil with reserves of methane and coal exceeding that of crude oil. It should then come as no surprise that with increasing crude oil prices, the inaccessibility of crude oil deposits and transportation cost of fuel to its markets, there is a shift toward the utilization of coal, natural gas and other alternatives sources for energy. A logical alternative to refining crude oil is the production of syngas (CO and H 2 ) from methane, coal or even biomass and furthermore, its conversion to a range of liquid fuels and chemicals through Fischer Tropsch Synthesis (FTS). Advantages of fuels derived from FT processes include the absence of sulfur, nitrogen or heavy metal contaminants, and a low aromatic content. With the focus of this work on JP-5 for the Navy; JP-5 produced from FT has good combustion properties and high smoke points. Potentially useful co-products are also generated such as waste heat which can be converted to electrical power and or reused back in the FT process. 2 Chapter 2 of this thesis gives a general overview of the fundamental concepts surrounding FT. The focus covers current FT catalyst, reactor types/design and operating conditions, the polymerization reaction and kinetics governing this reaction. Chapter 3 focuses on the choice of an appropriate catalyst for JP-5 production with the use of micro- fibers materials to enhance thermal conductivity. An optimal design for the catalyst is highlighted with respect to product selectivity which requires a comprehensive understanding of influences with promoters and supports. In addition, phase changes and catalyst deactivation are reviewed. The thermal conductivity measurement of a packed bed was compared to that of a micro-fiber catalyst bed showing significant improvement for heat removal. Chapter 4 details the use of this technology for specific case study examples with emphasis on problem formulation and significant results on JP-5 production. The impact of an improved FTS selectivity and the overall balance of plant (BOP) is assessed which includes a unique process design approach using simulation tools for upstream and downstream operations. Chapter 5 highlights the conclusions reached from the use of this technology, framework and iterative design, as well as a detailed plan of action to strengthen and expand this research area. It will be demonstrated that this research provides a method to address the increasing energy demands and crisis we face including those by the DoD. 3 ?The era of procrastination, of half-measures, of soothing and baffling expedients, of delays, is coming to a close. In its place we are entering a period of consequences?? Sir Winston Churchill, November 1936 Chapter 2 Theoretical Background 2.1 History In 1902, Sabatier and Senderens discovered that methane could be formed from a mixture of H 2 and CO over nickel and cobalt catalyst. German chemical manufacturer, BASF, produced liquid products such as paraffins, olefins and oxygenates on activated Co- Os catalyst at high pressure. 1 In time, nickel was abandoned as a catalyst because its activity declined due to the formation of nickel carbonyl. Haber and Bosch developed a method for ammonia synthesis in 1908 while Bergius worked directly on the hydrogenation of coal in 1913. 2 Franz Fischer and Hans Tropsch were responsible for the development of the Fischer Tropsch method in 1920 with the intention of producing high-value hydrocarbon molecules from coal-derived gas. Fischer and Tropsch successfully proposed the synthol process from CO and H 2 on alkalized iron chips. 3 In 1935, the first FT industrial application was designed; the Ruhrchemie atmospheric fixed-bed reactor. 4 By 1935, about 4 commercial-size, Ruhrchemie-licensed FT plants were under construction with an estimated annual capacity of producing 725,000 ? 868,000 barrels of gasoline, diesel fuel, and lubricating oil. 4 By 1944, 1 Dry, M. ?The Fischer Tropsch Process: 1950-2000.? Catalysis Today 71 (2002): p227-41 2 Guettel, R. ?Reactors for Fischer Tropsch Synthesis.? Chem. Eng. Tech. 31 (2008): p746-54 3 Khodakov, A. ?Advances in the Development of Novel Cobalt Fischer Tropsch for Synthesis of Long Chain Hydrocarbon and Clear fuels.? Chemical Reviews 107 N5 (2007): p1692-1744 4 Steynberg, A. ?Introduction to Fischer Tropsch Technology.? Studies in Surface Science and Catalysis 152 (2004): p1-63 4 another catalyst predominantly used in FTS was cobalt with a composition in relative mass units of 100Co:5ThO2:8MgO:200Kieselguhr (silicious diatomatous earth). 5 In 1950, the first FT plant was built in the United States in Brownsville, TX with the capacity to produce 5,000 bpd, but with an increase in feed prices, the plant was shut down. 2 In 2006, Sasol-Chevron embarked on a project to design a Gas-to-Liquid (GTL) plant in Escravos, Nigeria with an estimated production capacity of about 100,000 bpd. Exxon-Mobil signed an agreement with Al-Kaleen Gas Project (AKG) to build 150,000 bpd GTL-FT plants in Qatar. FT technology has reached an industrial full-scale application which has been commercialized worldwide. Table 1 below shows a timeline of FTS plant construction and the average capacity under which they operate. Table 1. Existing plants and plants under construction for Fischer-Tropsch synthesis. 2 Company Site Capacity Raw Material Commission Date [bpd] Existing plants Sasol Sasolburg 2500 Coal 1955 Sasol Secunda 85000 Coal 1980 Sasol Secunda 85000 Coal 1982 MossGas Mossel Bay 30000 Natural Gas 1992 Shell Bintulu 12500 Natural Gas 1993 Sasol/Qatar Petroleum Qatar 34000 Natural Gas 2006 Under Construction Sasol-Chevron Escravos Natural Gas 2007 Exxon-Mobil Qatar Natural Gas 2009 2.2 FTS Catalyst The four metals currently known to be most active in the hydrogenation of CO to hydrocarbon are Fe, Ni, Co and Ru. While Ni is the most active of the four metals, it is 5 Dry, M. ?The Fischer Tropsch Synthesis.? Catalysis: Science and Technology 1 (1981): p159-225 5 known to have a very high selectivity for methane while olefin selectivity is very low. 6 Additionally, the use of nickel as a catalyst creates inherent problems relating to the volatility of nickel carbonyls formed in the FT environment. Ru, Co, and Fe are therefore the catalysts of choice, however, the determination of the most desirable catalyst is still difficult to determine as several factors come into play including cost, availability, desired product spectrum, catalytic life time, and activity. Alkali-promoted iron catalysts have been industrially utilized in FTS for many years. Iron catalysts have high water-gas shift activity and selectivity toward olefins. 7 There are two types of iron catalyst, fused and precipitated iron. Fused iron catalysts are enhanced with an alkali to increase catalyst activity and selectivity. Precipitated iron catalysts, on the other hand, are doped with metals like Cu to enhance the reduction of iron oxide. The catalyst is then precipitated from an acidic solution with the addition of a basic solution, e.g. sodium carbonate or ammonia. Cobalt catalysts on the other hand, are usually supported with an alkali. A major difference between a cobalt catalyst and that of iron is that cobalt catalysts are not inhibited by water and thus result in a higher productivity at high syngas conversion. They are also known to give the highest yields and longest lifetime producing linear alkanes. 8 Cobalt disadvantages include a low water-gas shift activity and relatively expensive when compared to an iron catalyst. 6 Dry, M. ?FT synthesis over iron catalyst.? Catalysis Letters 7 (1990): p241-252 7 Jager, B. ?Advances in low temperature Fischer-Tropsch synthesis.? Catalysis Today 23 (1995): p17-28 8 Chaumette, P. ?Higher alcohol and paraffin synthesis on cobalt based catalyst: comparisons of mechanistic aspects.? Topics in Catalysis 2 (1995): p117-126 6 2.2.1 Product Selectivity The products from the FT synthesis form a complex, multi-component mixture varying in carbon number and product type. The product distribution of hydrocarbons can be described by the Anderson-Schulz-Flory (ASF) 9 equation; 1 )1( ? ?= n n m ?? (1) Thus, one is able to express this product distribution into weight fraction () 12 1 ? ?= n n nW ?? (2) () ? ? ? 2 1 logloglog ? += n n W n (3) The products of FTS over Co, Fe and Ru are shown on Figure 1. Figure 1 Hydrocarbon Selectivity on Co/TiO 2 200 o C; H 2 /CO ~ 2.1, and P ~ 20 bars. 10 Ru/SiO 2- 212 o C, H 2 /CO~2, and P~ 5 bars 18 Fused and precipitated Fe/Cu/K 11 9 Anderson, R. ?Catalyst for the Fischer-Tropsch Synthesis.? Van Nostrand Reinhold 4 (1956) New York 10 Iglesia, E. ?Selectivity Control and Catalyst Design in the Fischer-Tropsch Synthesis: Sites, Pellet, and Reactors.? Advances in Catalysis 39 (1993): p221-298 11 Donnelly, T.J. ?Analysis and Prediction of Product Distributions of the Fischer-Tropsch Synthesis.? Energy & Fuels 2 (1988): p734-739 7 where tp p kk k + =? ; and p k represent the propagation rate and t k represent the termination rate. 12 A summary of catalyst and corresponding alpha values are reported in Table 2. Table 2. Summary of Catalyst, Alpha values and Operating Conditions. 13 Element Alpha Operation Temperature (?C) Ru 0.85-0.95 180-250 Co 0.70-0.80 180-250 Fe 0.50- 0.70 220-250 (LT) 280-350 (HT) LT/HT ? Low/High temperature There are significant deviations from the ASF due to the uncertainties of not knowing the mechanism on the molecular level. These deviations have been well documented. For example; relatively high yields of methane, which have been explained by reactor hot spots, 13 increase termination of C 1 precursors, 14 low yields of ethane and ethene which is attributed to secondary reaction. 15 Changes in the growth parameter and the exponential decrease of the olefin to paraffin ratio with increasing carbon number have also been reported for most catalysts; iron 16,17 , cobalt 6,18 , and ruthenium. 8,19 This phenomena is attributed to the 12 Lox, E. ?Kinetics of the Fischer-Tropsch Reaction on a Precipitated Promoted Iron Catalyst, 2 Kinetics Modeling.? Ind. Eng. Chem. Res. 32 (1993): p71-82 13 Dry, M. ?Catalytic aspects of industrial Fischer-Tropsch synthesis.? J. Mol. Catal. 17 (1982): p133-144 14 Sarup B. ?Studies of the Fischer-Tropsch Synthesis on a cobalt catalyst. I. Evaluation of product distribution parameters from experimental data.? Canadian Journal Chemical Engr. 66 (1988): p831-843 15 Iglesia, E. ?Transport-enhanced ? -olefin re-adsorption pathways in ruthenium-catalyzed hydrocarbon synthesis.? Journal of Catal.129 (1991): p238-256 16 Rao, V. ?Iron-based catalysts for slurry-phase Fischer-Tropsch process: Technology.? Fuel Process technology 30 (1992): p83-107 17 Dictor, R. ?Fischer-Tropsch synthesis over reduced and unreduced iron oxide catalysts.? Journal of Catal. 97 (1986): p121-136 18 Madon, R. ?Selectivity in Catalysis? (S.L. Suib and M.E Davis, eds.), American Chemical Society, Washington, DC, (1993): p382 19 Inoue, M. ?Simple criteria to differentiate a two-site model from a distributed-site model for Fischer-Tropsch synthesis.? Journal of Catal,105 (1987): p266-269 8 occurrence of catalytic sites. Most of these deviations can be attributed to secondary reactions of ?-olefins. 20 2.3 FTS Reactors Pichler discovered in 1936 that when pressure was increased from atmospheric to about 15 bar, the life of iron catalyst improved significantly. 21 The original German industrial rector of 1936 (Lamellenofen) had 10 m 3 of packed catalyst between 600 parallel metal plates spaced 7 mm apart with 600 tubes. 22 After the 2 nd World War, Ruhrchemie and Lurgi formed Arbeitsgemeisnschaft (ARGE) and developed the fixed-bed reactor using a precipitated iron catalyst to produce high yields of wax. 5 The heat of the reaction was removed using water circulated through horizontal tubes which passed through the vertical plates. Due to low linear velocities of the gas, 100 hr -1 , the heat removal was not sufficient. This resulted in carbon deposition and disintegration of catalyst particles attributed to localized over-heating 3 . The next improvement was the development of the concentric tube reactor which had 2000 pairs of tubes with ID?s of 21 and 24 mm and 4.5m long. Cooling water was circulated outside of outer tube and through the inside of the inner tube but still low space velocity was still an issue. BASF utilized the concept of high recycle ratio (100 to 1) in a single bed reactor of catalyst (3.8 m 3 ) to attempt to resolve the temperature rise within the FTS reactor. 23 However, overheating still occurred leading Lurgi to improve such a technique by splitting the bed into several sections with fresh feed split, i.e. feed between successive beds. The overall recycle to fresh feed gas ratio per bed was 20 to 1 and the temperature rise per bed was about 5 o C. 4 The reactor was operated at 270 o C, 20 bar and a 20 Iglesia, E. ?Design, synthesis, and use of cobalt-based Fischer-Tropsch synthesis catalyst.? Applied Catalysis A 161(1997): p59-78 21 Pichler, H. :ibid (1952) vol. 4 22 Frohning, C. ?Fischer-Tropsch Synthese.? Chemierohstoffe aus kohle (Falbe J., ed.) (1977). Stuttagart: Thieme 23 Ullmanns Encyklopadie d. techn. Chem. 9, 715 Munchen-Berlin: Urban and Schwarzenberg (1957) 9 fresh feed space velocity of about 200 hr -1 . A high conversion of 85% was obtained. 3 The use of a high gas linear velocity through the catalyst bed ensures that heat generated is removed along the length of the tubes and this results in a nearly isothermal reactor at the expense of compression cost. 2.3.1 Reactor Design There are four types of FT reactor commercially used today, - Tubular fixed-bed reactor - Circulating fluidized-bed reactor - Fluidized bed reactor - Slurry phase reactor A picture of these reactors is shown in Figure 2. Figure 2 Types of Fischer-Tropsch Synthesis Reactors. 24 Because of the strong exothermic nature of FT reaction, the control of temperature within the reactor is the most important issue for a safe and stable operation. Other important 24 Steynberg, A. ?Fischer Tropsch Reactors.? Studies in Surface Science and Catalysis 152 (2004): p64-195 10 reasons for monitoring the temperature within the reactor include life of catalyst and product selectivity. In this regard, the industrial preference is to use a slurry reactor over fixed bed because of its high heat transfer characteristics. 25 The slurry reactor operate within the range of 320 o C to 350 o C, hence the term high temperature Fischer Tropsch (HTFT) while that of the tubular fixed bed and slurry phase reactor operate within the 220 o C to 250 o C range and have been termed low temperature Fischer Tropsch (LTFT). Beyond the operating temperature range that differentiates between the HTFT and LTFT reactors is the fact that there isn?t a liquid phase present outside the catalyst particles in the HTFT reactors. This is because of particle agglomeration and loss of fluidization. 26 0 10 20 30 40 50 CH4 C2?4?Olefins C2?4? Parraffins Gasoline Distillate Oils?and? Waxes Oxygenates Produc t ? Di s t r i b u t i o n ? (p e r ? 10 0 ? ca rb o n ? at o m s ) LTFT?200?? 250C HTFT?270?? 350C Figure 3 Selectivity of FTS for LTFT and HTFT on Fe Catalyst. 27 25 Song, H. ?Operating Strategies for Fischer-Tropsch Reactors: A Model-Directed Study.? Journal of Chemical Engineering 21(2004): p308-317. 26 Baird, M. ?FT Process investigated at the Pittsburgh Energy Technology Center since 1944.? Ind. Eng. Chem. Prod. Res. Dev., 19 (1980): p175-191 27 Mulder, H. ?From Syngas to clean fuels and chemicals via Fischer-Tropsch process.? Presented at Gasification, the gateway to a cleaner future, Dresden Germany, Sept. 1998. 11 2.3.2 Fixed Bed Reactors Fixed bed reactors are vertically spaced packed bed and radial flow reactors with cooling tubes between the beds. The multi-tubular reactor is packed with catalyst inside the tube walls (by using narrow tube diameters) and usually operates at high gas linear velocities to improve the transfer of the heat generated from the catalyst to the cooling medium. The use of narrow tubes ensures that the distance between the hot catalyst particles and heat exchanger surface is short and also ensures that the ratio of heat exchanger surface area to catalyst mass is high. There are radial as well as axial temperature gradients in these reactor tubes, about 2 o C to 4 o C (tube centre to wall) and an axial temperature difference of 15 o C to 20 o C (peak in upper section to exit) 28 with the high linear space velocities. There are advantages to using a fixed bed reactor; these are; easy to operate. They can be used over wide range temperature ranges irrespective of whether the FT products are gaseous or liquid, or both, under reaction conditions. There are no problems with separating liquids products from the catalyst. However, fixed bed reactors are prone to pressure drop, however this can be reduced by working with a high catalyst bed voidage of about 90%. 29 Some more disadvantages include cost, expensive to construct. The FT reaction is diffusion controlled and so it would be advantageous to use small catalyst particles. Replacing the catalyst is labor intensive due to the narrow tubes. Typical fixed bed reactors operate at low temperature, but with iron catalyst a high temperature, 260 o C, with an operating pressure of 28 bar and a fresh space velocity of about 1000 hr -1 (recycle ratio of about 20) is commonly used. 30 28 Dry, M. ?Practical and theoretical aspects of the catalytic Fischer-Tropsch Process.? Applied Catalysis A 138 (1996): p319-344. 29 Forney, A. ?US Bur. Mines Rep.? 5841ACS Div. Fuel 20 (1961) 30 Fischer, F Roelen, O.:ibid.11,489 (1930) 12 2.3.3 Fluidized Bed An alternative way of rapidly transferring heat is to move the particles to the heat exchanger which is what actually occurs in a two phase fluidized bed. The high gas velocity limits film diffusion at both the particle surfaces and the heat exchanger surfaces. The direct physical contact between hot catalyst particles and heat exchanger tube walls also contributes to heat exchange. The combination of these factors results in a much higher efficiency of heat exchange within the fluidized bed which means a smaller heat exchanger area is required. Another advantage of fluidized beds is that once catalyst particles are fluidized, the differential pressure across the reactor will not increase with further increase in gas velocity. The pressure drop is proportional to the mass of catalyst being used. Commercial fluidized beds operate in turbulent regimes and this result in a nearly isothermal reaction zone with differential temperatures across the bed of 2 o C or less. Fluidized beds do have a few disadvantages owing to their operational complexity. Separation of the catalyst from the exhaust gas is not simple. There is erosion of the tubes and fouling due to the abrasive iron carbide catalyst. This adds to cost and increases reactor down time. 31 Two-phase fluidized systems cannot be used when wax is to be produced and typical fluidized bed reactors operate at temperatures of 300 o C and above. 2.3.4 Slurry Bed Reactors Gas can be bubbled through a suspension of finely divided catalyst particles (typically 10 -200?m) in a liquid which has a low vapor pressure. The heat generated is removed by circulating the slurry through external heat exchangers or immersing heat exchangers directly 31 Jager, B. ?Experience with a new type of reactor for Fischer-Tropsch synthesis.? Catalysis letters 7 (1990): p293-301 13 into the slurry bed. The influence of internal mass transfer resistance is negligible and both optimal activity and selectivity are achieved. The Slurry reactor is a fixed fluidized bed. In FT applications, it is used for the production of high molecular waxes which are liquids under synthesis conditions and also why the liquid of choice is usually a cut from the product spectrum, e.g. high boiling wax. A slurry bed has a direct advantage over a fixed bed in the sense that it can be used easily for HTFS operations. Comparing this reactor to a ?dry? fluidized bed, it has the advantage of operating with lower H 2 /CO ratio. Disadvantages of the slurry bed include scalability and the separation of the solid catalyst. Advantages and disadvantages of both established reactor technologies are summarized in the table below. Table 3. Advantages (+) and disadvantages (-) of established reactors for FTS. 2 Fixed bed reactor Bubble column Reactor Pore diffusion - + Catalyst content in reactor + - Gas-liquid mass transfer + - Isothermal behavior - + Catalyst exchange - + Catalyst attrition + - Need for liquid-solid separation + - Scale up + - Reactor cost - + (+) Advantages and (-) Disadvantages Both reactors technologies exhibit certain disadvantages: the fixed-bed multi-tubular reactor suffers from high pressure drop, low catalyst utilization and insufficient heat removal, whereas slurry bubble column reactor faces the need for catalyst separation, less ideal residence time behavior and highly demanding scale up. An ideal reactor would have the following characteristics: - Fixed bed catalyst - High catalyst efficiency due to short diffusion distances - Highly efficient gas-liquid mass transfer 14 - Isothermal operation at the highest possible temperature 2.4 FT Polymerization Reaction The overall reactions for an FT process are listed below: 32 Main reactions Paraffins (2n+1)H 2 + nCO ? C n H 2n+2 + nH 2 O Olefins 2nH 2 + nCO ? C n H 2n +nH 2 O WGS reaction CO + H 2 O ? CO 2 + H 2 Side reactions Alcohols 2nH 2 + nCO ? C n H 2n+2 O+ (n-1)H 2 O Catalyst Ox/Red (a) M x O y +yH 2 ? y H 2 O + xM (b) M x O y +yCO ? y CO 2 + xM Bulk carbide formation yC + xM ? M x C y Boudouard reaction 2CO ? C + CO 2 The FT synthesis product spectrum consists of a complex multi-component mixture of linear and branched hydrocarbons and oxygenated products. The main products are linear paraffins and alpha-olefins. 2.5 FT Mechanism/Kinetics FTS kinetics has been studied extensively and many describe the rate using a power law model or certain mechanistic assumptions that are thermodynamically controlled. These empirical equations that describe the intrinsic rate of Fischer-Tropsch carbon monoxide hydrogenation have been proposed using a mechanism that highlights the rate of formation of monomer as the rate limiting step. Anderson et al. postulated a method to describe the 32 Van der Laan, G. ?Kinetics and Selectivity of the Fischer-Tropsch Synthesis: A Literature Review.? Catal. Rev. ? Science Engineering 41 (1999): p255-315 15 product formation by successive addition of C 1 units to growing chains on a catalyst surface. 33 He also postulated that the rate of consumption of carbon monoxide for cobalt catalyst was proportional to the rate of desorption of the hydrocarbon chains growing on the catalyst surface. 9 There are many general theories that have been postulated. Carbide theory was the first and describes how CO dissociates and forms surface metal carbide. The intermediate reacts to form a methlyene group which is said to polymerize into a hydrocarbon chain. ***2 1 OCCO K +???+ ** 22 2 +????+ OHHO K ** 22 CHHC k ???+ The model derived from this theory describes how CO dissociates on the surface; adsorbed carbon reacts with hydrogen in the rate-determining step. This leads to a rate reaction of OHHCO CO FT PPPKkK PPKkK r H 22 2 21 2 21 1+ =? (4) The Enolic theory is the second, where it is believed that adsorbed CO is hydrogenated to a hydroxylated species. The chain growth forms from the elimination of water via condensation. ** ' 1 COCO K ????+ ** 2 ' 2 2 +????+ OHOH K ** 2 ' 2 COHHCO k ???+ This leads to an equation of 33 Anderson, R. ?Fischer?Tropsch Reaction Mechanism Involving Stepwise Growth of Carbon Chain? J. of Chem. Phys.19 (1951): p313 16 COOH HCO FT PKKP PPKKk r )/'( )'/'(' 21 21 2 2 + =? (5) Steen proves extensively that this is flawed because methane is the most abundant product and the polymerization character of FTS along with observed product distribution are never accounted for in the popularly postulated mechanisms. 34 The rate equation proposed by van Steen is based on the assumption that the rate of the reaction in FTS is governed by the rate of hydrogenation of surface carbon. The formation of organic compounds in FTS features a polymerization reaction with three classes of reactions: initiation, propagation and termination. Steen defines the reaction initiation as the formation of a chain starter from the reactants hydrogen and carbon monoxide. The propagation step is the incorporation of monomer units into growing chains which is the chain growth step in FTS. The monomer produced in situ on the catalyst surface. The termination is desorption of growing chains from the catalyst surface. The rate of carbon monoxide consumption is equal to the sum of the initiation rate, propagation and termination steps. Carbon is not involved in the termination reaction therefore the rate of consumption of carbon monoxide equals the rate of initiation plus propagation. This is also the rate of formation of chain starts plus the rate of incorporation of carbon into the growing chain (chain growth). In FTS the rate of incorporation is a function of the carbon number of the growing chain. 34 Steen, van Eric. ?Polymerization Kinetics of the Fischer-Tropsch CO hydrogenation using iron and cobalt based Catalyst.? Applied Catalysis A 186 (1999): p309 17 Figure 4 kinetic scheme of the successive hydrogenation of surface carbon. 34 At steady state, the net rate of formation of CH 2 surface species equals zero. Therefore, the rate of formation of CH 2 surface species by hydrogenation of CH species ( 2 CHCH r ? ) equals the rate of incorporation. The rate of formation of CH 2 surface species is thus equal to the rate of carbon monoxide consumption for the formation of organic compounds. 22 3, onincoporatiCHCHCHorgC rrrr +== ? The rate of consumption of monomer units is equal to the sum of the rates of incorporation into chains of all possible lengths. The net rate of formation of CH 2 surface species equals the net rate of formation of CH species. The rate of consumption of carbon monoxide for the formation of organic compounds is expressed as the rate of formation of CH surface species. 2 , CHCHCHCorgC rrr ?? == Because the formation of CH surface species can be regarded as irreversible, the net rate of formation of surface species is proportional to the fractional coverage of surface carbon and the fractional coverage of surface hydrogen; HCCHCCHC kr ?? ?? = The rate of formation of organic compounds on a carbon basis can principally be derived by using steady state approximation of C, H, CO, O and OH surface species. For simplicity, the species are in equilibrium with the gas phase compounds. 18 H 2 + 2* ? 2H* * 5.0 2 ? ? H H H P K = CO + * ? CO* * ? ? CO CO CO P K = CO* + * ? C* + O* * ?? ?? CO OC C K = O* +2H * ? H 2 O + 3* H O OH OH p K 2 3 * 2 2 ?? ? = The postulated equilibrium is assumed to occupy approximately the same geometric space. The rate of consumption of carbon monoxide can be now be expressed as 2 * 2 3 3 , 2 2 2 ? OH H OHCCOHHorgC p p KKKKCkcr ?= With a reasonable assumption that the surface is mainly covered with surface carbon under reaction conditions: 1 * =+ C ?? Which gives the rate of organic compounds as () 22 2 3 3 , )/(1 )/( 222 2 2 2 OHCOHOHCCOH OHCO H OHCCOHH orgC pppKKKK pppKKKKCkc r + ? = [] 2 2 3 , )/(1 )( 222 2 OHCOHOH CO H orgC pppbp ppa r + = (6) 2.5.2 Water-Gas Shift Reaction The water-gas shift can increase or decrease the rate of FTS. This is because of the shared components of adsorption and desorption reactions as well as dissociation of H 2 , H 2 O 19 and CO 2 . There are a few studies that have reported empirical kinetic expressions for the WGS reaction under FTS conditions using iron catalysts; 35 )( 22 2 2 P HCO COOHwCO K pP pPkr ?= (7) Because WGS is an equilibrium reaction, 36 close at or close to under FTS conditions, the reverse reaction has to be taken into account; )029.2 2073 (log ?= T Kp (8) It is believed that the reaction of WGS may take place on the same catalytic sites as those in FT. 36 The authors were not able to derive reliable kinetics equations for WGS under FT conditions and the question remains as to how WGS competes with FTS. Taking full advantage of the above technology concept of combining the length- independent chain growth process with a selective chain length-dependent conversion process is a technology that will help meet the growing energy demand. Though there is still a need to reduce the investment required and improve operating efficiency in commercial FT process applications, GTL fuels offer very attractive options to the hydrocarbon industry to introduce clean fuels. The FT process can and will be tuned so as to yield JP-5 in the next chapters. 35 Huff, G. ?Intrinsic Kinetics of the Fischer-Tropsch Synthesis on a Reduced Fused-Magnetite Catalyst.? Ind. Eng. Chem. Process Des. Dev. 23 (1984): p696-705 36 Graaf, G. ?Chemical equilibria in methanol synthesis.? Chem. Eng. Sci. 41 (1986 ): p2883-2890 20 "Civilization is in a race between education and catastrophe. Let us learn the truth and spread it as far and wide as our circumstances allow. For the truth is the greatest weapon we have." H.G Wells. Chapter 3 FTS Catalyst Design for Jet Fuel Abstract This chapter focuses on the design of an optimal catalyst matrix to enhance JP-5 selectivity with the aid of a highly thermally conductive microfibrous material. A promoted iron catalyst, Fe/Cu/K-La- Al 2 O 3 , was ultimately chosen because of its high alpha value of 0.87 and high water gas shift activity. The effects of promoters and supporters, investigated by pioneers and their relationship towards iron catalysts, are highlighted in this chapter. Kinetics of different catalysts was used to calculate the catalyst bulk activity with promoted iron showing high promising values. In addition, the thermal conductivity for a packed bed was measured and compared with that of a microfibrous bed. The microfibrous material increased the effective bed thermal conductivity by 90%. The ability to improve the bed conductivity will allow for a more prominent JP-5 selective control in FTS. 21 3.1 Catalyst Design and Optimization The interaction between the reaction kinetics and transport phenomena in a catalytic process takes place on a number of different levels; from the atomic site, to the catalyst size, and reactor operating conditions. Most FT studies and evaluations have separated catalyst development from reactor technology and development. A different design approach has been applied in this thesis, merging the two different paths, from the active site level of one?s catalyst to modeling scale up facilities, in an iterative manner. The relationship is symbiotic because active sites lie within porous pellets which control the rates of diffusive processes that supply reactants and remove products while the reactor allows for convective transfer of mass and heat over the catalyst bed. The added advantage of working simultaneously on these two steps for a catalyst design for FT is to achieve high metal particle dispersion and catalytic conversions. This will ensure that potential losses are minimized and the labor required to design an overall process for FT is reduced. In order to achieve optimum process design and product selectivity for Fischer Tropsch Synthesis (FTS), the two different paths of design that have in the past been separated must be coupled. Commercial catalyst screening was chosen as a logical starting point. After the new catalyst has been synthesized, its catalytic performance is tested and compared to that of conventional and commercial catalysts. This is known as catalyst screening. Catalyst synthesis, activations, pretreatments, evaluation of catalytic performance, and characterization are the next primary steps in the catalyst design. Comparisons of characterization data and results of catalyst testing allows for the nature of active sites to be characterized and catalyst synthetic routes to be optimized. 22 Figure 1 A Integrated Approach in FTS Catalyst Design Framework and Process Design for Balance of Plant. 1 The next stage of the catalyst design is the generation of catalytic systems whose structure, composition, actives sites and catalytic performance have been qualitatively defined by characterization and screening as well as the development of a molecular simulation. Kinetic studies then provide a quantitative relationship between the intrinsic reaction rates and composition of fluid around the catalyst. Developing the intrinsic kinetic model of the catalytic reaction will lead to the next step which would be linked to molecular simulations and thus modeling and evaluation of a pilot reactor testing. JP-5 is composed of C 9 -C 16 , linear hydrocarbons with no aromatics. It should come as no surprise that to design such a catalyst for FTS, the first step should focus on the catalyst, ?, chain growth probability value and, by employing the Anderson-Schulz-Flory 2 distribution; such values should be around 0.86 to achieve the maximum C 9 -C 16 selectivity. 1 Khodakov, A. ?Advances in the Development of Novel Cobalt Fischer Tropsch for Synthesis of Long Chain Hydrocarbon and Clear fuels.? Chemical Reviews 107 N5 (2007): p1692-1744 2 Anderson, R. ?Catalyst for the Fischer-Tropsch Synthesis.? Van Nostrand Reinhold 4 (1956) New York 23 Figure 2 shows the weight distribution of FTS. The product selectivity can be achieved by optimizing the syngas ratio, operating conditions of the FT reactor and with the use of an efficient heat removal technique with an appropriate catalyst that has been promoted. 0 0.2 0.4 0.6 0.45 0.55 0.65 0.75 0.85 0.95 W e i g ht Fr a c t i o n Chain Growth Probability C2?C4 C1 C5?C8 (Gasoline) C9?C16 (JP?5) C17+(Waxes) Figure 2 ASF Weight distributions as a function of chain growth probability The objective of this work is to design an appropriate catalyst structure that will achieve an optimum performance with the aid of a microfibrous material. Microfibrous materials provide a mechanical and electrical entrapment of a particle/fibrous solid such as sorbents or catalysts within a sinter-locked network of a secondary fibrous matrix such as a metal or polymer fiber (MFEC). With MFEC, there are certain advantages over a conventional packed bed reactor or a reactor packed with monolithic catalyst structures such as high heat transfer removal rates due to a high thermal conductive reactor bed. 3 3 Kalluri, R. ?Microfibrous Supported Sorbents/Catalysts - Micro-Structured Systems with Enhanced Contacting Efficiency.? Presented at AIChE Annual Meeting. (2005) 24 3.2 Catalyst Matrix Ruthenium is a very active catalyst and studies have reported its ability to have high alpha values in FTS. 4 However, it is relatively expensive as a catalyst when compared to cobalt and iron catalysts, and consequently commercial reactors typically avoid this catalyst. 5 In addition, ruthenium is highly selective toward producing large amounts of methane at low pressures and at low temperatures; a shift towards high amounts of waxes is seen. 6 Co and Ru catalysts promoted with Zr, Ti or Cr are widely used to for the production of waxes which can be cracked to diesel fuels and olefins. 5 Conventional catalysts used to make various premium products are summarized in Table 1. Table 1. Summary of Premium Products, Catalyst and Technology Used. 7 Premium product Catalyst Reactors Gasoline Fused Fe/K Fluid-bed Co/ThO 2 /Al 2 O 3 /Silicalite Fixed-bed Fe/K/ZSM-5, Co/ZSM-5 Slurry/Fixed-bed Ru/ZSM-5 Fe/Cu/K and ZSM-5 Diesel Fuel Fe/K, Ru/V/TiO 2 Fixed-bed (low T) Co/Zr, Ti or Cr/Al 2 O 3 Slurry bed (low T) Co/Zr/TiO 2 Co-Ru/Al 2 O 3 Waxes Fe/K, Fe/Fe/Cu/K Slurry-bed (low T) Co/Zr, Ti or Cr/Al 2 O 3 Fixed bed (low T) Co/R/Al 2 O 3 , Fe/Ru The three key properties for FT catalysts that many have optimized are product selectivity, activity and lifetime. Each one of these properties can be affected by a variety of 4 Dry, M. ?Catalytic aspects of industrial Fischer-Tropsch synthesis.? J. Mol. Catal. 17 (1982): p133-144 5 Rao, V. ?Iron-based catalysts for slurry-phase Fischer-Tropsch process: Technology.? Fuel Process technology 30 (1992): p83-107 6 Dry, M. ?The Fischer Tropsch Synthesis.? Catalysis: Science and Technology v1 (1981): p159-225 7 Bartholomew, C. H. ?Recent technological developments in Fischer-Tropsch catalysis.? Catalysis Letters 7 (1990): p303-314 25 strategies including the use of promoters, catalyst preparation and formulation, pretreatment and reduction, and shape selectivity. Early precipitated iron catalysts were promoted with potassium salts and thus, to improve the stability of these catalysts towards thermal sintering, supports like zinc oxide, kieselguhr, and alumina were added. The influence of varying SiO 2 on the total area and pore structure of Fe 2 O 3 is apparent in the following figures, figure 3 and 4. It is believed that SiO 2 increases the area and stabilizes the presence of larger pores. 6 Figure 3 Influence of SiO 2 on pore volume of Fe 2 O 3 . 6 . Figure 4 Influence of SiO 2 on Surface Area of Fe 2 O 3 . 6 26 There are a large number of different recipes used for the preparation of heterogeneous catalysts for FT. The inherent compositional and structural inhomogenity of these supported systems make selecting the right catalyst extremely difficult. Preparation, techniques involving impregnations, controlled pH precipitations, and decomposition of metal carbonyl on dehydrated supports all affect the activity of the catalyst, and thus alpha value during FT. 8,9 The activity of a catalyst and the product selectivity are also influenced by the reducing temperature with which the catalyst is reduced. When a catalyst is reduced in H 2 at 220 o C for example, it allows for higher molecules such as waxes to be formed as opposed to another catalyst when reduced at 280 o C. 9 While preparation techniques are important in the determination of the pore structure of a catalyst and its activity, it is not the focus of this research. 3.2.1 Catalyst for Jet Fuel Production Fe/Cu/K has been chosen as the desired catalyst for JP-5 production because of its related high activity as a catalyst, its ability to achieve high alpha values, and subsequently high fractions of C 9 -C 16 paraffin and waxes. In addition, a high water-gas shift activity makes a promoted iron also a favorable catalyst for FTS. The effects of potassium promoters on iron catalysts have being well documented by pioneers such as: Anderson et. 1952 10 /1956; 11 Dry and Oosthuizen, 1968, 12 Bonzel and Krebs, 1981, 13 Dry 1981, 4 Arakawa and Bell, 1983, 14 8 Bukur, D. ?Pretreatment effect studies with a precipitated iron Fischer-Tropsch catalyst.? Applied Catalysis A 126 (1995): p85-113 9 O?Brien, R.J and Davis, B.H. ?Activation of Precipitated Iron Fischer Tropsch Catalyst.? Energy & Fuels 10 (1996) : p921-926 10 Anderson, R. ?Fischer-Tropsch Synthesis, Some important variables of the synthesis on iron catalyst.? Ind. Eng. Chem. 44 (1952): p391-397 11 Anderson, R. ?Catalyst for the Fischer-Tropsch Synthesis.? Van Nostrand Reinhold 4 (1956) New York 12 Dry, M. ?Heat of Chemisorptions on Promoted Iron Surfaces and the Role of Alkali in Fischer-Tropsch Synthesis.? Journal of Catalysis 15 (1969): p190-199 27 Dictor and Bell, 1986, 15 Donnelly and Satterfield, 1989. 16 With respect to promoting iron with potassium and copper, Bukur studied the effects of these specific promoters on iron?s alpha values, product selectivity as well as catalyst activity for different catalyst compositions in a fixed bed at temperatures of 235-365 o C and 15 bar (Bukur, 1990). 17 Bukur reported that both potassium and copper increased the catalytic alpha distribution and activity for FTS and water gas shift, suppressing the secondary reactions (olefin hydrogenation and isomerization of 1-alkenes to 2-alkenes). The reported alpha values for the two different regimes normally seen in FT 18 are shown in Table 2. Table 2. Chain Growth Probability factors for different Catalyst 17 . Catalyst?? 1 ?? 2 ? 100?Fe? 0.56???0.11?? 100?Fe/0.2?K? 0.61???0.08? 0.84???0.04? 100?Fe/0.5?K? 0.68???0.02? 0.94???0.03? 100?Fe/1?K? 0.66???0.02? 0.88???0.02? 100?Fe/3?Cu? 0.63???0.03? 0.91???0.03? 100?Fe/3?Cu/0.2?K? 0.62???0.09? 0.79???0.01? 100?Fe/3?Cu/0.5?K? 0.69???0.02? 0.93???0.06? Three different catalysts were tested and evaluated, Fe/Cu/K/La-Al 2 O 3 , the second as Fe/Al 2 O 3 and Fe/SiO 2 /Al 2 O 3 were tested under FT conditions at 20 bars and H 2 /CO ratio of 13 Bonzel, H. P. ?Enhanced Rate of Carbon Deposition during Fischer-Tropsch Synthesis on K Promoted Fe.? Surface Science Letters 109 (1981): pL527-531 14 Arakawa, H. ?Effect of potassium promotion on the activity and selectivity of iron Fischer-Tropsch Catalyst.? Ing. Eng. Chem. Process Des. Dev. 22 (1983): p97-103 15 Dictor, R. ?Fischer-Tropsch synthesis over reduced and unreduced iron oxide catalysts.? Journal of Catal. 97 (1986): p121-136 16 Donnelly, T. J. ?Product Distributions of the Fischer-Tropsch Synthesis on Precipitated Iron Catalysts.? Applied Catalysis 52 (1989 ): p93-114 17 Bukur, D. ?Promoter Effects on Precipitated Iron Catalyst for Fischer Tropsch Synthesis.? Ind. Eng. Chem. Res. 29 (1990): p194-204 18 Huff, G. ?Evidence for two chain growth probabilities on Iron Catalysts in Fischer Tropsch Synthesis.? Journal of Catalysis 85 (1984) p370-379 28 2. Figure 5 was constructed from the collected data and the calculated values of alpha from the product distribution with respect to operating temperature. 19 Figure 5 Alpha for Fe/Cu/K, Fe/Al 2 O 3 and Fe/ Al 2 O 2 /SiO 2 (20bars and H 2 /CO ~2) With increasing hydrogen concentration and temperature, an empirical equation was developed for ?; 20,21 one that takes into account both the H 2 /CO ratio and temperature relationship. This means that the potential product distribution could be obtained with respect to temperature and concentration at any point in the reactor. )]260(0039.01[ 2 ??+ + = TB yy y A HCO CO ? (1) Tables 3 show the values of A and B that were obtained from a Non-Linear regression analysis on the data in Figure 5. Figure 6 show the influence of Syngas ratio on alpha. 19 IntraMicron, Inc/Cerametec Topic N07-T027, Proposal Number N2-2578 (2009) 20 Yermakova, A. ?Thermodynamica Calculations in the Modeling of Multiphase Processes and Reactors.? Ind. Eng. Chem. Res. 39 (2000): p1453-1472 21 Lox, E. ?Kinetics of the Fischer Tropsch Reaction on a Precipitated Promoted Iron Catalyst, 2. Kinetic Modeling.? Ing. Eng. Chem. Res. 32 (1991): p71-82 29 Table 3 Estimated constants A and B Catalyst A B Fe/Cu/K/La-Al 2 O 3 0.984 ? 0.34 0.46 ? 0.11 Fe/Al 2 O 3 -3.41 ? 2.46 1.59 ? 0.72 Fe/SiO 2 /Al 2 O 3 -1.29 ? 2.32 0.91 ? 0.69 See Appendix V for non linear analysis. Figure 6 Alpha as a function of H 2 /CO on Fe/Cu/K/La- Al 2 O 3 Catalyst. A low-temperature Fischer-Tropsch (LTFT) reaction would be appropriate for achieving this high alpha and JP5 selectivity. There is also the advantage of operating within this region to avoid the formation of fouling ?coke?. LTFT would extend the life of a catalyst as opposed to operating within a high-temperature Fischer Tropsch (HTFT) regime and dealing with the issues of associated phase changes of the Iron catalyst. 7 In addition, at high temperatures, the Boudouard reaction becomes more prominent 22 , resulting in a continuous increase in elemental carbon content at the surface of the catalyst. Carbon deposition becomes apparent on the catalyst, directly affecting the decline of the catalyst and fouling the tubes. 22 Dry, M. ?The Fischer-Tropsch (FT) synthesis processes.? Handbook of heterogeneous catalysis. 6 (2008): p2965-2994 30 Kolbel et al. 23 concluded that potassium when used as a promoter stabilizes the surface area of precipitated iron oxy-hydrides and protects against re-crystallization during calcinations. In addition, there were more active sites found than in an un-promoted catalyst. However, using higher concentrations of potassium as a promoter will cover active sites, resulting in a decrease in catalyst activity. While there are no proven theories, an increase in the overall activity of a catalyst in the presence of copper has been observed. Dry reports that copper facilitates the reduction of iron thus decreasing the total surface area of the catalyst. 7 The effects of the promoters, supports, additives, pretreatments and preparation are only meaningful in the comparison of baseline catalytic properties in the absence of these effects. In order to calculate the activity (or turnover frequency), a proposed reaction mechanism is needed. A few related mechanisms have been reported as well as the associated rate formula derived for such mechanisms. Eliason reported rate data and proposed a power law expression based on such data for unsupported Fe and Fe/K catalysts which was then used to calculate a bulk activity (Eliason and Bartholomew, 1999 24 ). Atwood (Atwood and Bennett, 1979 25 ) reported intrinsic kinetics over a commercial promoted fused iron catalyst, Fe/Al 2 O 3 /K 2 O 3 /SiO 2 . Liu (Liu and Li, 1995 26 ), and Leib (Leib and Kuo, 1984 27 ) reported Langmuir-Hinshelwood (LHHW) expression for our proposed catalyst, Fe/Cu/K, which was derived through the Carbide theory. 23 Knobel, H. ?Beitrag zur Fischer-tropsch Synthese on Eisenkontakten.? Chem. Ing. Tech. 23 (1951): p153 24 Eliason, S. A. ?Reaction and deactivation kinetics for Fischer-Tropsch synthesis on unpromoted and potassium-promoted iron catalysts.? App. Catalysis A 186 (1999): p229-243 25 Atwood, H. ?Kinetics of the Fischer-Tropsch reaction over iron.? Ind. Eng. Chem. Process Des. Dev. 18 (1979): p163-170 26 Liu, Z. ?Intrinsic kinetics of Fischer-Tropsch Synthesis over Fe-Cu-K Catalyst.? J. Chem. Soc. Faraday Trans. 91 (1995): p3255-3261 27 Leib, T. ?Modeling the Fischer-Tropsch Synthesis in Slurry Bubble-Column Reactors.? Paper presented at the AIChE Annual Meet., San Francisco, California (Nov. 25-30, 1984) 31 Table 4. Summary of FT Kinetics and Parameter Values from Literature. Source Catalyst Rate T(C) Ko a Atwood and Bennett (1979) CCI fused Iron 250 0.0017 (mols/gcat-s-Mpa) E (kJ/mol) 0.028 85 -8.8 Leib and Kuo (1984) Fe/Cu/K 265 0.062 (mols/gcat-s-Mpa) E (kJ/mol) 0.58 85 -8.8 Liu and Li (1995) Fe/Cu/K 210 7.83E-6 (mols/cm3-s-Mpa) E (kJ/mol) 66.5 1.9E-5 E (kJ/mol) 65.4 Eliason and Bartholomew (1999) Unpromoted Fe 265 3.77E+6 (mols/kg-Cat-s-atm ^0.98) E (kJ/mol) 101 Fe/K 265 3.77E+6 (mols/kg-Cat-s-atm ^0.55) E (kJ/mol) 92 Turnover frequency, is the number of revolutions of the catalytic cycle per unit time, generally second. 28 As a measure of the activity of a catalyst, turnover frequency is the number of molecules of CO converted per catalytic site per second. This has been calculated for different catalyst structures for FTS operating temperature 210 o C (LHFT), 260 o C (HTFT) and pressure of 20 bars. The turnover frequencies were calculated based on the assumption of 10% dispersion and a constant loading. Figures 7 and 8 show this comparison between the estimated turnover frequencies of promoted iron catalyst. 28 Boudart, M. ?Kinetics of Heterogeneous Catalytic Reactions? Princeton University Press: Princeton, NJ, 1984 32 Figure 7 CO hydrogenation activities As seen in Figure 8, Fe/Cu/K gives a high turnover frequency. At high temperatures a higher extent of activity is seen because these rates occur faster. Figure 8 CO hydrogenation activities. 4 33 In a study conducted by Jam, 29 Fe/Cu/La/SiO 2 was compared with Co/Al 2 O 3 . Cobalt was revealed to have a smaller particle diameter size. The table below shows average particle size and BET surface area of the catalyst. Table 5. Characterization data of Catalyst 29 . Catalyst? Bet?Specific?Area?(m 2 /g)? Particle?Diameter?(XRD)? Co? 154? 14.4(nm)? Fe? 47? 60(um)? Figure 9 CO hydrogenation activities 5 It can be inferred from figure 9 and table 5 above that at a cobalt catalyst would be a better choice for FTS to make JP-5 given its high activity and surface area, making it easier to achieve higher per pass conversions over an iron catalyst. However, more research needs to be conducted on the use of cobalt as a catalyst because it is difficult to obtain consistent 29 Jam, S. ?Enhancement of Distillate Selectivity in Fischer Tropsch Synthesis by Using Iron and Cobalt Catalyst in a Novel-Dual Bed Reactor.? Reaction Kinetics and Catalysis Letters 89( 2006): p71-79 34 literature data for the rate conversion of CO. Many reports highlight certain short comings for cobalt catalysts, for example, its rapid decline in activity due to the buildup of long chain waxes in the pores. 30,31 It has also been found that for a Co/Al 2 O 3 catalyst, the presence of water in the syngas results in surface oxidation of the cobalt, causing permanent loss in activity. 32,33 In addition, Kogelbauer and coworkers report that with Co/silica, an increase in amount of silicates is formed with time under FT conditions, causing cobalt to lose its activity. 34 In spite of these reports, however, defining a catalyst activity is subjective. Under proper conditions and depending on operating regimes, both catalysts have the potential to operate six months or longer. 35 The following table summarizes a desired catalytic function for FT reactions. Table 6. Optimization of Fe Catalyst Design. Desired Catalytic Functions Proven Catalyst Components/ Structural Features Key Aspects of Preparation and Pretreatment High Structure Integrity with MFEC Dense alumina, silica and titania carriers stabilized with Ba, La, or Zr oxides. Modify Al 2 O 3 surface with promoter like La Extent of Reduction addition of Cu, Pt, Pd, or Ru; Dope Fe precursors with promoted metal, K, Pt, Pd High resistance to fouling by carbon and heavy waxes Relatively large mesopores (d pore = 12-16 nm); additives such as Cu, Ru and Pt which gasify carbon Choose ??Al 2 O 3 with d pore = 12-16 nm; uniform dispersion of additives like Cu 30 Van Berge, P.J. ?Cobalt as an alternative Fischer-Tropsch catalyst to iron for the production of middle distillates.? Studies in Surface Science and Catalysis 107 (1997): p207-212 31 Niemela, M.K ?The long-term performance of Co/SiO2 catalysts in CO hydrogenation.? Catal. Letters (1996) v42, p161-166 32 Schanke, D. ?Study of the deactivation mechanism of Al2O3-supported cobalt Fischer-Tropsch catalysts? Catalysis Letters 34 (1995):, p269-284 33 Hilmen, A. ?Study of the effect of water on alumina supported cobalt Fischer-Tropsch catalysts. Applied Catalysis. A 186 (1999), v186, p169 -188 34 Kogelbauer, A. ?The formation of cobalt silicates on Co/SiO2 under hydrothermal conditions.? Catalysis. Letters. (1995) v34 p259-267 35 Davis, B.H. ?Fischer-Tropsch Synthesis: Comparison of Performances of Iron and Cobalt Catalysts.? Ind. Eng. Chem. Res. 47 (2007): p8938-8945 35 High activity High C 9-16 selectivity High surface area High Fe loading and site density; large mesopores (d pore = 12-16 nm); High dispersion (10-15%); add Cu, or Ru to enhance Fe reduction; add K to increase C 9+ selectivity Inert support; uniform catalyst distribution; drying, calcination and reduction at low heating rate and high SV; uniform dispersion of additives 3.2.2 Phase Changes during Synthesis There can be many changes within the phase composition when using an iron catalyst. This is due to the ease with which iron can be oxidized as well as the formation of iron carbides. The metallic iron phase is not stable under normal FT conditions. Below is a plot of the phase diagram of iron magnetite with time during an FT operation. Figure 10 The change in composition of iron catalyst during FT reaction 300C. 6 Figure 10 represents an illustration of a possible phase change with time at 300 o C for an iron catalyst in FT. The phases are those which are present according to X-ray diffraction analysis. At time equals zero, the catalyst is 100% metallic iron. The units are undefined since the rates of the phase changes depend on the amount of alkali present. A phase diagram is shown in Appendix III for iron magnetite with respect to temperature and pressure. 36 3.2.3 Influence of Promoters on Physical Properties and Synthesis Fe catalysts are significantly affected by the chemical nature of promoters and/or supports. The promotion with an optimal amount of alkali materials, such as potassium salts, affects the FT activity as well as the required selectivity. 6 The first concern is the influence of promoters on the physical properties of iron. Using pure iron oxide as the host, the influence of various cations has been studied extensively. Al 3+ , Li + and Mg 2+ have a decreasing effect in the Fe 3 O 4 unit cell as the promoters are increased. For larger cations like Ca 2+ , Mn 2+ , Ti 3+ and Na + , the unit cell size increases as these promoters are increased. 36 K + ions do not have any effect on the unit cell because they are too large to replace the Fe ions in Fe 3 O 4 . Arakawa et. al 13 in his studies showed that potassium promotion on an alumina supported iron catalyst causes a decrease in the Fe dispersion, however, it increased the strength of CO chemisorptions on reduced Fe. The presence of the promoters inhibits crystal growth of the ?-Fe and thus results in a high surface area of the reduced catalyst. 37 This promotional effect has been described by Dry and it is due the ratio of the ionic charge to ionic radius of the cations. Therefore, the more ?acidic? the cation, the larger the surface area of your catalyst. In a nutshell, Al 2 O 3 and TiO 2 strongly increase the area while CaO and MnO have little effect. 36 Dry, M. ?The distribution of promoters in magnetite catalysts.? Journal of Catalysis 7 (1967): p352-358 37 Dry, M. ?The influence of structural promoters on the surface properties of reduced magnetite catalysts.? Journal of Catalysis 6 (1966): p194-199 37 Figure 11 Influence on the unit cell size of magnetite (in Am) of various promoters in solid solution with magnetite. 6 See Figure 11 and 12 for the influence of the unit cell size of magnetite and typical results of the influence of promoters on area. If the catalyst were operated at high temperatures, the associated reaction rates appear to be diffusion controlled and thus an increase in internal surface area will show some benefit. Figure 12 The surface area of fully reduced catalyst in m2 (g unreduced sample)-1 as function of promoter content (g atom promoter cation per 100 g atom Fe). 6 38 It is important to promote the iron catalyst with strong alkali metals of Group 1, (i.e. K 2 O). This is because precipitated catalysts are unstable and thus need to have their available active surface area increased and stabilized to a level required for high conversions during FTS. The higher the level of alkali, the greater the relative activity as shown in Table 7. Table 7. Influence of alkali content on synthesis performance of precipitated iron catalyst 6 . Unsupported Fe 2 O 3 Silica Supported Fe 2 O 3 K 2 O level a Activity a K 2 O level a Activity a 0 26 12 112 1 47 16 109 1.6 50 21 85 2 53 24 83 3 40 32 75 a ?relative?quantities 3.2.4 Catalyst Size and Shape The effect of the catalyst size can be seen in Figure 14. CO conversion increases when particle sized is reduced from 0.3mm to 0.154mm, highlighting the point that intra- particle diffusion is eliminated when particle size is reduced. This is a simulation study that was obtained using Aspen TM . It can be verified experimentally. Figure 13 Effect of particle size of the catalyst on the conversion of CO (Conditions: T = 210C, 20bars, H/CO = 2, L = 100mm GHSV = 600hr -1 ) 39 The effictiveness facor can be calculated; certain assumption such as a pseudo first- order reaction rate, k=0.06s -1 and diffusivity of H 2 . 38 Utilization of smaller catalyst particles improves selectivity toward larger molecular weight hydrocarbons promoting intra-particle mass transport. This will allow reeducation of the catalyst loading while maintaining the same or higher volumetric reactivity. Higher effectiveness factors and smaller particles provide greater surface to volume ratios and better activity maintenance. These are represented in Figure 13 and 14. Figure 14 Catalyst Effectiveness Factor Table 8. Some common catalyst shape. 39 Shape Comments Spheres Low manufacturing cost. Some granulated material can be weak, and cementing agents are frequently used to increase strength. Packed beds of spheres produce relative high pressure drop Irregular Granules Not a common catalyst shape. Consequence of restricted methods of manufacture. Used for low surface area catalyst such as bulk metal 38 Post, M.F.M. ?Diffusion Limitation in Fischer-Tropsch Catalyst.? AIChE Journal 35 (1989): p1107-1114 39 Twigg, M. V ?The Catalyst: Preparation, Properties and Behavior in Use.? Catalysis and Chemical Process (1981) R. Pearce and W.R Patterson p11 40 Pellets Highly regular shape, good strength. Although expensive to produce, the most common catalyst shape. Extrudate Low bulk density and low pressure drop. Often poor strength. Some formulations can be extruded which are difficult to pellet. Extrusion aids (additives) and cementing agents are frequently used. Production cost is much less than pelleting Rings Impregnation of preformed rings can produce very high strength catalyst. Low pressure drop, and so used in tubular reactors. Manufactured by pressing or extrusion Pressed rings have superior physical properties but costly to produce. 3.3 Thermal Conductivity Thermal conductivity is key to the idea of using a metal micro-fibrous entrapped catalyst to improve the rate of heat transfer. The adiabatic temperature rise in FTS is about 1750 o C, and conventional reactors for fixed beds have use recycle gas/diluents such as nitrogen with high space velocity, monolithic catalyst structures 40 or a slurry reactor to solve this problem. Since weight and volume is a constraint for this operation, it is important to limit the size of this reactor, therefore the use of a bubble column is not an option. It is imperative that heat removal is performed efficiently to avoid a loss in product selectivity. To determine the thermal conductivity of the packed bed for comparison with an MFEC bed, a steady-state temperature profile inside a heated tube with flow through a homogeneous material was measured on a packed bed of alumina powder and a micro- fibrous copper sample. A segregated 2-D heat balance for the gas, catalyst and fiber phases were written to model the tube heat up. PDEs describing the heat transfer inside the bed were discretized using central finite differences with radial averaging. These equations 40 Mesheryakov, V.D. ?A multifunctional reactor with a regular catalyst packing for Fischer-Tropsch synthesis.? Chemical Engineering Science 54(1999): p1565-1570 41 describing the transient heat transfer were integrated into steady-state using a line implicit Euler method in the axial direction. The simulation was programmed using the Euphoria interpreter. The derivation of the segregated 2-D model is given in the Appendix II. Two materials were tested for this comparison. One with a bed of 80x100 mesh alumina powder. The other with a sintered mesh consisting of equal amounts of 4 and 12 micron copper fibers and the same alumina powder. A two inch section of a 1 ?? O.D. stainless steel tube was filled with the media to be evaluated. The tube was sealed with five thermocouples inside during the first set of tests. Figure 15 Diagram of thermocouple locations in the packed tube heat-up test. In a second set of heat-up tests, eleven thermocouples were used. A diagram of the location of the thermocouples in the packed bed heat-up test is shown in Figure 15 above. More results of this work, including temperature profiles and calculations can be seen in a work which is currently being developed. 41 The temperature profile obtained and a fitted profile is shown in Figure 16 and 17. 41 Sheng M. ?Effective Thermal Conductivity of Metal Microfiber Entrapped Catalyst.? (2010) Auburn University Unpublished 42 Figure 16. Temperature Bed comparisons between Packed Bed and MFEC (Copper fibers and alumina powder-80x100). 40 50 60 70 80 90 100 0 0.25 0.5 0.75 1 Ce n t e r l i n e ?Te m p . ?/C e l c i u s z/L Steady?State?Temperature?Profile?for?Copper?MFEC 23slpm 23slpm 18slpm 18slpm 12slpm 12slpm 6slpm 6slpm Figure 17. Segregated 2-D simulation result for heat-up of a tube packed with MFEC (Copper fibers and alumina powder-80x100). Table 9. Thermal Conductivity and Wall heat transfer coefficients Packed bed MFEC k(W/mK) 0.16 ? 0.35 9.05 ? 0.21 h(W/m 2 K) 22.7 ? 10.2 235?30.5 Adding mass balances and heat generation in the catalyst phase to a segregated 2-D model would be a simulation of the FTS reaction with uniform velocity at any cross section. This would be a further development on the work already done and will allow the 43 determination of an effective thermal conductivity measurement during FT conditions. There are radial and axial temperature profiles that need to be considered also. These profiles could produce a limitation in the performance of iron kinetics because of the fixed bed characteristics. All possible scenarios should be explored. Conclusion The ideal Fe catalyst will be supported on a relatively inert support. It?s ability to be well reduced with the addition of a noble metal makes such a catalyst very good choice for FTS. Such a catalyst will contain Cu/K as a promoter that enhances Fe reduction and limits carbon formation. Hypothetical compositions include Fe/Cu/K/Al 2 O 3 or Fe/Pt/K/La 2 O 3 /Al 2 O 3 . With the use of a micro-fibrous entrapped catalyst, intra-bed heat transfer can be enhanced which will able to maintain the temperature profile within the reactor. This in turn will improve the product selectivity for JP-5 within the FT reactor. 44 ?It isn?t that they can?t see the solution. It is that they can?t see the problem.? G.K. Chesterson Chapter 4 FTS Process Design for Jet Fuel Abstract The feasibility of using Micro-Fibrous Entrapped Catalysts (MFEC) in Fischer-Tropsch Synthesis (FTS) for selective JP-5 production as well as overall balance of plant reductions (BOP) will be addressed in this chapter. Three different simulation case studies based on a minimized weight by volume analysis of a Gas-To-Liquid (GTL) process under optimum reaction conditions of, 210 o C and 20 bars, were carried out using an iron promoted catalyst, (Fe/Cu/K) entrapped in 8 vol.% copper microfibers for the benefit of a mobile skid unit. The use of MFEC significantly enhanced the selectivity to JP-5 (33wt%), reduced the balance of plant (BOP), 20 to 30% and improved the utilization of natural gas (14%) while maintaining the same production capacity when compared to a conventional GTL plant with a packed bed reactor. By greatly reducing the size of process equipment, this in turn will reduce the cost; these are achievements that are attractive for a GTL offshore plant design. 45 4.1 Introduction GTL is a promising method of converting syngas to fuel products. The conversions of natural gas to synthetic fuel has attracted significant attention due to current economic uncertainties surrounding crude oil prices, low margins on crude refining and, more importantly the need to become less dependent on foreign crude oil supplies. 1 Currently, the major source of energy in the United States is petroleum, as seen in Figure 1; with 37% of this source being crude oil and, 20% of crude oil imported from Canada. 2 Coal 23% Natural?Gas 24% Petroleum 37% Nuclear?Electric? Power 9% Renewable? Energy?(Biomass) 4% Geothermal?Energy 0% Hydroelectric? Conventional 2% Sola/PV?Energy 0% Wind? Energy 1% Figure 1 U.S Energy Consumption. 2 Natural gas without access to the world market is known as ?stranded gas.? This includes large reserves in remote places and associated gas that is co-produced with crude- oil. There are an estimated 900 to 3,000 trillion cubic feet in volume of available natural gas; which includes stranded gas. 3 Natural gas is either co-produced with petroleum or rests on 1 Speight, J. ?Liquid Fuels from Natural Gas.? Handbook of Alternative Fuel Technologies. (2007): p153-170 2 U.S Energy Information Administration ? Independent Statistics and Analysis. < http://tonto.eia.doe.gov/energy_in_brief/foreign_oil_dependence.cfm> May 2010. 3 Petroconsultants-MIA and Zeus International ?Remote Gas Development Strategies.? HIS Energy Services (1999) 46 top of petroleum reservoirs. With no local market for this gas, oil production is not possible without venting, flaring, or reinjection of the gas back into the reservoir. Venting is not an option as the global warming potential of methane is about 21 times that of CO 2 . Gas flaring is also not desirable because of stringent environmental policies in effect in a large number of countries. The current cost of reinjection is about $13 per equivalent barrel, with offshore gas reinjection being more expensive, thus making this option economically infeasible. 3 GTL synthetic fuels play a significant role in the environment as their use results in low emissions of CO, and NO x . 4 Although coal is another major source of energy and its role in Fischer-Tropsch Synthesis (FTS) is resurging in the United States, 5 it has a number of disadvantages to overcome such as considerable amounts of CO 2 , SO x , NO x and pollution- causing particulate matter (PM) being formed. Coal-to-Liquid (CTL) technology will be a vital source of energy supply when an acceptable and efficient way of CO 2 sequestration is found. Synthetic fuels produced from FTS range from light gases, which fall into the category of LPG to middle distillates, which comprise of mainly C 7 -C 22 hydrocarbon chains such as jet fuel, kerosene, and diesel, to the heavies, C 20+ , known as ?wax.? Figure 2 shows a descriptive comparison of products obtained from an FTS process and those produced via conventional crude oil refineries. 4 Li, X. ?Particle size distribution from a GTL engine.? Sci. Total Environment.382 (2007): p295-303 5 Eilers, J. ?The Shell Middle Distillate Synthesis Process (SMDS).? Catalysis Letters 7 (1990): p253-269 47 Figure 2 Comparison of FT Products with conventional Refinery barrel. 6 The United States Department of Defense (DoD) specifically the Navy, has several reasons for an interest in JP-5 fuel produced from alternative energy resources and processes, such as FTS. 7 Most of these reasons are similar to those discussed earlier with others worth mentioning; supply chain vulnerability and fuel supply continuity. In the case of a natural disaster, there is a concern related to acquiring fuel because of the heavily reliance upon crude oil imports, and the majority of crude-oil refineries are located on the East, Gulf and West coast, which are vulnerable to attack. Another attractive reason is that FTS technology has the ability to manufacture fuel near or within remote locations not easily accessible for pipelines to be built. This would seem vital as there is a high cost associated with transporting fuel to battlefield endured by the Military. GTL plants can be mounted on barges or ships in order to be moved from one site to another as needed. This enables monetization of many smaller fields where reserves 6 Ronald, M. ?New developments in Gas to Liquid Technology.? CERI 2004 Petrochemical Conference, Alberta, Canada. (2004). 7 Forest, C. ?Fischer-Tropsch Fuels: Why are they of interest to the United States Military?? SAE Tech. Pap. Ser. (2005) Refinery Barrel (vol.%) GTL - FT Barrel (vol.%) 48 are smaller or rapid depletion is indicated, as constructing a gas pipeline would present a large financial risk. Many drivers for use of any technology offshore are safety, size, reliability and cost. When one considers each of these factors, a floating or air based GTL technology is suitable for offshore use. The DoD currently uses a single kerosene-type fuel for all its gas-turbine and diesel applications which comply with Jet-Propulsion 8 (JP-8), specifications. JP-8 is similar to commercial aviation turbine fuel but with required additives such as corrosion inhibitors, icing inhibitors and static dissipator additives. 8 JP-8 is a kerosene type fuel whereas diesel fuels are generally a distillate blend, a distillate and kerosene blend, or a kerosene blend depending on the grade of diesel which differ from state to state. There are additives in JP-8 such as lubrication improvers, biocides, antioxidants, and thermal stability improvers are also used in specified amounts as governed by military specifications. 9 JP-5 is used on aircraft carriers which is similar to JP-8 but with a higher flash point of 60 o C. A higher flash point provides an additional degree of safety for handling fuels. 10 JP-5 is a complex mixture of hydrocarbons and naphthalene?s consisting of C 9 -C 16 linear alkanes. Synthetic jet fuels produced from alternative energy sources have to comply with the MIL-DTL-5624T specification. This specifies the fuel?s physical and chemical properties such as freezing point, flash point, density, heating value, hydrogen content and smoke point. 11 A comparison between the specifications of diesel fuel and FTS JP-5, as JP-5 is the focus of this work, is seen in Table 1. 8 Muzzell, P. ?Composition of Syntroleum S-5 and Conformance to JP5 Specification.? Prepr. Pap.-Am. Chem Soc., Div., Pet. Chem. 49 (2004): p411-413 9 DOD JP-5 and JP-8(1992). May 2010 10 Lamrecht, D. ?Fischer Tropsch Fuel for Use by the U.S. Military as Battlefield-use Fuel of the Future.? Energy & Fuels 21 (2007): p1448-1453 11 MIL-DTL-5624U Detail Specification, Turbine Fuel, Aviation, Grades JP-4 and JP-5 < http://www.desc.dla.mil/DCM/Files/5624t.pdf> May 2010 49 Table 1. Selected Fuel Properties for Jet Fuel and Diesel Units MIL-DTL-5624T (JP-5) 11 Diesel (ASTM D975) 12 S-5 (JP-5 from FTS) 13 Density at 15?C kg/L 0.788 - 0.845 0.810 - 0.95 0.767 IBP 10% o C 206 ~130 182 FBP o C 300 360 195 Flash Point o C 60 52 62 Viscosity cSt 8.5@-20 o C 4.1 @38 o C 7.0 @-20 o C Freezing Point o C -46 <0 <50 Heating Value MJ/kg 42.6 43-49 44.1 Cetane No. - 40 62 Smoke Point mm 19 - >43 Aromaticity % vol., max 25 35 - Before GTL conversion can become an investment of the future, certain technological challenges must be addressed, specifically for a mobile skid unit. The theoretical maximum thermal efficiency for the conversion of coal to paraffin is 60%, while that of natural gas is approximately 75%. The efficiency attainable in practice is lower, which is why proper heat integration must be performed so that heat generated onsite is reused in the process while minimizing external cooling and heating requirements. Furthermore, capital investment for installation is a limiting factor; with reports of Sasol?s capital cost for a 34,000 bpd costing $25,000 bpd. 14 The hydrocarbon condensates and wax produced by an FTS process are usually seen as value added fuels, however, it will require additional energy and capital investment to hydrogenate the olefins and oxygenates formed as well as the use of a hydro- cracking unit in order to break down higher chain molecules into shorter chains. Other challenges facing GTL facilities involve syngas production: justifying the economy of scale when choosing among steam methane reforming (SMR), partial-oxidation (POX) and auto- 12 Diesel Fuel Spec. Sheet ASTM D975 < http://xtremefueltreatmentdenver.com/ASTM_D975_TEST.pdf> May 2010 13 Freerks, R. ?Production and Characterization of Synthetic Jet fuel Produced From Fischer Tropsch Hydrocarbons.? Prepr. Pap.-Am. Chem. Soc., Div. Pet. Chem. 49 (2004): p407 14 Thacheray, F. ?GTL in 2007.? Petroleum Reviews (2003) p18-19 50 thermal reforming (ATR). The DoD also must be concerned with introducing FTS fuels into its military fleet which currently runs on crude-oil derived fuels which contain aromatics. There is a fundamental need for applied innovative research into catalytic materials as well as structural architectures that enable FTS to be transformed into a modular, flexible and inexpensive technology. Due to the exothermic nature of FTS, fully packaged catalyst structures are required which possess high inherent heat transfer rates offering the opportunity for a better control on intra-bed hot spots and product selectivity. A high selectivity towards JP-5 would reduce the BOP requirements for post processing operations. High inherent heat transfer rates also permit fundamental changes and simplifications in the form factor and geometry of the FTS reactor with respect to the number of individual reactor tubes. The design of this plant should be easily transported in modular fashion by ship and operated with quick startup properties; these are traits that are highly desirable for the Navy. The objective of this chapter is to assess the possibilities of reducing the BOP by developing and utilizing a process simulation analysis on a steady state performance with the use of micro-fibrous entrapped catalyst (MFEC) in a fixed bed reactor with the goal of producing JP-5. The use of an MFEC reactor will not be subjected to drawbacks that conventional fixed bed reactors face, (e.g. hot spots and high gas compression cost). Process operating conditions such as temperature and reactant partial pressures as well as syngas composition are addressed in order to enhance the selectivity of JP-5. A preliminary Phase 1 engineering design and evaluation is performed on the possibilities of making the FTS reactor smaller and thus its effect on the overall BOP. An analysis based on weight and volume for the skid unit was evaluated for 3 different operations. 51 4.2 Integrated Simulated Designs Aspen Plus TM was used as the process simulator with material and energy balances solved for all process units. It was assumed that the system operates at steady-state and natural gas feed was held constant. A heterogeneous model was used to model this problem. The formation of olefins and alcohols were ignored for simplifications. The physical properties of reaction medium were calculated by the Peng-Robinson equation of state. The NTRL equation of state was applied for a three phase system; a situation where separations of liquid-liquid-vapor mixture was needed. The following compounds have been selected from Aspens databank: O 2 , N 2 , CO, CO 2 , H 2 , and paraffins from CH 4 -C 30 H 62 with a minimum number of other heavy hydrocarbons introduced into Aspen because of its lack of property definitions for higher chain alkanes greater than C 31+ . The properties added for higher chain alkanes were vapor pressure, density, molecular weight and boiling point from API and ASTM tables. Linear and saturated hydrocarbons are selected to describe gasoline (C 5 -C 11 ), JP-5 (C 9 -C 16 ) and waxes (C 20+ ). A plug flow model is used to model the FTS reaction in Aspen using a promoted iron catalyst, Fe/Cu/K on alumina. A flow scheme of a typical FTS system is presented in figure 3. Reformer/ Gasifier Gas Cleanup/ Conditioning Co or Fe Slurry /Tubular Reactor Fe Fluidized Reactor Hydrocracking Oligomerization Isom erization Hydorogenation Coal, Natural Gas Air, Oxygen, Steam LTFT HTFT Waxes(>C 20 ) - Gasoline/Diesel - Jet Fuel - Gasoline/Diesel - Jet Fuel olefins Figure 3 Flow Scheme of an FTS Process 52 An FTS system consists of at least 4 stages; upstream, (where syngas is produced and purified), Fischer-Tropsch Synthesis (FTS) reactor, separations of products, hydro-cracking, and water/product treatment and recycling. 4.2.1 Upstream The manner in which syngas is produced is influenced by many factors, which in turn impacts many aspects of the rest of a GTL design. Some of these factors include plant size and location, the need for an oxygen plant, gas compression, heat integration and gas recycle options as well as configuration of power generation alternatives. One can calculate the thermal efficiencies from the syngas producing facilities through the ratio of enthalpy of syngas to that of the feed (methane and/or oxygen) used, as written in Equation 1. Tables 2 and 3 compare syngas generating facilities to one another. pyFeedEnthal alpySyngasEnth iciencyThermalEff =% (1) Table 2. Comparison of Syngas generation technologies (Natural Gas Feed) I 15 Technology Advantages Disadvantages SMR - Most extensive Industrial experience - H 2 /CO ratio often higher than required when CO also is to be produced - Oxygen not required - Highest air emissions - Lowest process temperature requirement - Highest air emissions - Best H 2 /CO ratio for hydrogen production applications POX - Feedstock desulfization not required - Can be a disadvantage when requiring H 2 /CO ratio >2 - Absence of catalyst permits carbon formation and therefore, operation without steam significantly lowering syngas CO 2 content - Very high process operating temperatures - Low methane slip - Usually requires Oxygen 15 Tindal, A. ?Natural gas reforming technologies.? Gas-to-Liquids Processing; Bringing Clean Fuels to market, San Antonio, TX, March 18-20 1998 53 - Low natural H 2 /CO ratio is an advantage for applications - High temperature heat recovery and soot formation/handling adds process complexity - Low natural H 2 /CO ratio is an advantage for applications - Syngas methane content is inherently low and not easily modified to meet downstream processing requirements ATR - Natural H 2 /CO ratio often is favorable Limited commercial experience - Lower process temperature requirement than POX - Usually requires oxygen - Low methane slip - Syngas methane content can be tailored by adjusting reformer outlet temperature Table 3. Comparison of Syngas generation technologies II. SMR POX ATR Operating Temperature, o C 900 1500 1100 Pressure, bars 25 25 25 Thermal Efficiency, % 75 79 82 CO/H 2 Ratio Normal 3.5 2 2.5 Recycle CO 2 1.9 1.6 1.6 Increase Steam >5.0 >1.8 >2.65 Auto-thermal (ATR) proves to be more thermally efficient due to the fact that it is a combination of partial oxidation (POX) with a catalytic steam reforming (SMR). However, SMR results in a higher production of hydrogen compared to carbon monoxide in the product syngas and therefore, will be the preferred technological route for this facility because of this very reason. Another attractive reason for choosing SMR as the favorite syngas producing technology is an oxygen plant is not required. This will be one less unit which minimizes the weight and volume design for a mobile GTL facility. It has been well advised that steam methane reforming by itself is not the preferred technology for syngas production for large 54 scales (>10,000bpd) GTL facilities because of the economics but this technology is sufficient for a 500 bpd plant. 16 Synthesis gas for low temperature Fischer Tropsch (LTFT) operations has an optimum H 2 /CO ratio slightly greater than 2 with low CO 2 concentration. 17 However, a hydrogen starved syngas increases chain growth probability, and thus selectivity hence, a target ratio of H 2 /CO of 1.9 from the SMR unit will be set (See Chapter 3 for effects of syngas composition on alpha). A flowsheet of the upstream process is shown in Figure 4. Figure 4 Process Flow Diagram of FTS - Upstream. The key reactions in steam methane reforming are; CH 4 + H 2 O ? CO + 3H 2 O (2) CH 4 + CO 2 ? 2CO + 2H 2 O (3) CO + H 2 O ? CO 2 + H 2 (4) It has been assumed most content of sulfur in the natural gas stream has been stripped, thus leaving a mole fraction composition of 1% N 2 , 96% CH 4 , 1.5% C 2 H 6 and 0.5% C 3+ (by mole fraction). 15 The furnace is modeled as two blocks in Aspen, with one being the reactor 16 Aasberg-Perterson, K. ?Synthesis gas production for FT synthesis.? Studies in surface science and catalysis. 152 (2004): p258 17 Lu, Yijun. ?Influence of the Feed Gas Composition on the Fischer-Tropsch Synthesis in Commercial Operations.? Journal of Natural Gas Chemistry 16 (2007): p329-341 55 (SMR) and the other furnace. The heat from the SMR is used to determine the heat duty of the furnace that is needed. Two design specs were used to obtain the correct amount of air and natural gas feed needed to set the outlet temperature of the SMR at 900 o C. The effluent gas from the furnace is used to heat various other streams in the plant before it is vented into the atmosphere. Water is converted to pressurized steam at 500 o C and natural gas (stream NGAS1) is preheated to 500 o C at 20 bars using the heat from the furnace. These streams heat up to just slightly above 500 o C, the minimum temperature suggested by IEA, 18 and react through the 900 o C reactor which is assumed to achieve equilibrium concentrations. One benefit of this design is that all of the required steam is produced within the plant. With CO 2 being recycled back into the SMR, we are able to achieve the desired hydrogen to carbon monoxide ratio of 1.9:1 with over 65% conversion of methane. The addition of recycled CO 2 , causes a water gas shift reaction to occur, reducing the ratio of H 2 /CO ratio for desirable LTFT conditions. A sensitivity analysis was performed by varying the steam/natural gas ratio and temperature in the feed stream to maximize the energy efficiency and H 2 /CO ratio of 1.9 of the SMR unit (results displayed in table 4). The variation of the steam to natural gas volume ratio and natural gas feed temperature are expected to improve the energy efficiency of the SMR. Decreasing the ratio of steam to natural gas with a constant CO 2 recycle favors the desired ratio of H 2 /CO but does not increase the thermal efficiency. This is can be attributed to the effect of the recycled CO 2 . The natural gas temperature was varied with a constant steam to volume ratio of 30%. High natural gas feed temperature increases the thermal efficiency of the SMR. 18 Precombusion Decarburization, IEA Greenhouse Gas Program, Report #PH2/19, p22 56 Table 4. Sensitivity Analysis of steam to methane ratio and N.G temperature on GTL Performance. Steam/N.G (vol.%) Thermal Efficiency % H 2 /CO ratio N.G temperature o C Thermal Efficiency % 20 73.48 1.71 400 73.60 30 73.64 1.92 510 76.25 50 75.01 2.36 600 79.42 After the syngas exits the steam reformer, it is cooled down to 70 o C in a heater. This heat is used to produce 9,000kg/hr of pressurized steam that can be used as feed, operate a turbine engine or, as indicated in this simulation, to operate steam heated bottoms reboiler for the distillation columns. The separation of CO 2 from the syngas stream can be achieved with the use of methanol via the Rectisol process as cold methanol absorbs carbon dioxide between 25-30 bars. The Rectisol process requires more electrical energy for refrigeration to maintain the low temperatures required (-44 o C) but also requires less steam energy for regeneration. 19,20 Another method for removing CO 2 is through the use of N-methyl- diethanolamine (MDEA), removal of CO 2 is achieved by condensation and drying and through a pressure swing adsorption unit, other contaminants can be removed. MDEA is preferred as it is less energy intensive and operates at a favorable temperature of 57 o C. 21 These processes are more complicated and require a complex model of their own for complete representation and would not be address in this thesis but for the purpose of the overall modeling effort, a black box model have been utilized for simplicity. After the syngas is cooled, the water is separated out in a flash in order to reduce reactor size and to prevent condensation. This is also performed in order to prevent a competing water-gas shift reaction that would reduce the selectivity of JP-5. Almost 1700kg/hr of water is being removed and reused as a coolant or raw feed. 19 Liu, J. ?The Rectisol process for natural gas purification.? Natural Gas Chemistry 32 (2007): p47-50 20 Kohl, A. ?Gas Purification? (5th ed.), Gulf Publishing, Houston, TX (1997). 21 Aliabad, Z ?Removal of CO 2 and H 2 S using Aqueous Alkanolamine Solution.? World Academy of Science, Engr. And Technology. 49 (2009): p194-203 57 4.2.2 FTS Reactor The FTS reaction is highly exothermic and has an adiabatic temperature rise of up to 1750 o C. The problem of heat removal which many are faced with is most efficiently solved by using a slurry reactor. This in practices provides isothermal operation. There are a number of advantages besides temperature control for the use of a slurry reactor: high extent of catalyst use, design simplicity, ability of on-line catalyst regeneration and lower operation expenses. There are a few distinct disadvantages associated with this type of reactor such as the need for separating the liquid-catalyst phase from the product stream and for scale up purposes, this unit will not be ideal for a mobile FTS skid unit because of its size. With a fixed bed reactor, heat removal from the catalyst particles is done conventionally through high gas velocities of the reaction mixture flow and with diluents such as nitrogen. This results in additional energy expenditures to overcome hydraulic resistance of the catalyst bed. Attempting to use large particles to reduce the hydraulic resistance of the catalyst bed will only result in a decrease in the effectiveness factor of the catalyst as the process then is controlled by intra-particle diffusion. There has been a recent growing interest in monolith reactors, 22 but there are problems associated with these as well, such as low volumetric reactivity and flow distribution. Micro-fibrous entrapped catalyst (MFEC) reactors combine the advantages of a fixed and a slurry reactor without their corresponding disadvantages assuming operating equipments are properly maintained. Table 5 summarizes the advantages and disadvantages of each of these reactors and highlights the benefits of a MFEC reactor. 22 Mesheryakov, V.D. ?A multifunctional reactor with a regular catalyst packing for Fischer-Tropsch synthesis.? Chemical Engineering Science 54(1999): p1565-1570 58 Table 5. MFEC Reactors versus Fixed and Slurry Reactors. Fixed bed reactor MFEC bed reactor Bubble Column reactor Pore diffusion - - + Catalyst content in reactor + + - Gas-liquid mass transfer + + - Isothermal behavior - + + Catalyst attrition + + - Need for liquid-solid separation + + - Scale up + + - Reactor cost - - + (+) Advantages and (-) Disadvantages In order to determine the overall heat transfer coefficient and model a bed temperature profile, one needs to be able to calculate the wall heat transfer coefficients; specifically for wall-to-fluid and wall-to-solids. Dixon?s detailed works on effective heat transfer in a packed bed 23 provides the theoretical background necessary to calculate these coefficients. The effective axial and radial thermal conductivities are known and as such, an effective intra-bed thermal conductivity and overall heat transfer coefficient can be calculated. The Nusselt number relationship for a packed bed is defined: ) PrRe 1()(4 RF fw p Pe Nu R D Nu ?? ++= (5) The wall heat transfer coefficient with respect to fluid and solid respectively; p rg wf d NuK h * = ; p rs ws d K h 12.2* = (6) The effective radial thermal conductivity is calculated and used in solving an overall heat transfer coefficient. 23 Dixon, A. ?Theoretical Prediction of Effective Heat Transfer Parameters in packed Beds.? AIChE Journal 25 (1979): p663-676 59 ) ))(1( ) 1.01 ( 3 16 1 8 1 ( 2 p t rspfs rs twf rg rsrgr d d kdh k dh k kkk ?? + + + += (7) The Overall Heat transfer coefficient 24 4 3 3 11 * + + += Bi Bi k R hU r t w (8) The relationship for concentration and temperature dependence in a 1D heterogeneous model are listed below; 25 AB A s r dz dC u ?=? (9) )(4)( w t ABpBs TT d U rH dz dT Cu ????= ?? (10) Pressure drop 19 () () ? ? ? ? ? ? ? ? ? + ? ?= ? ? 2 3 2 23 2 )( 1 75.1 )( 1 150 GG ppB B GG ppB B dd z P ?? ?? ? ?? ?? ? (11) More details can be found in Appendix IV. A recent kinetic study by Liu 26 allows for the use of intrinsic kinetics for FT synthesis over the design catalyst of choice, Fe-Cu-K. The lumped reaction kinetics are listed below. OHCO HCO HCO P RT P PP RT R 2 2 2 ) 412,65 exp(10*8973.1 ) 520,66 exp(10*2178.1 5 2 ? + + ? =? (12) The effectiveness factor on the Fe-Cu-K catalyst has been reported by Wen Jie 27 24 Dixon, A. ?An improved equation for the overall heat transfer coefficient in Packed beds.? Chemical Engineering and Processing 35 (1996): p323-331 25 Froment, G. ?Chemical Reactor Analysis and Design.? John Wiley & Sons (1979) New York 26 Liu, Z. ?Intrinsic kinetics of Fischer-Tropsch Synthesis over Fe-Cu-K Catalyst.? J. Chem. Soc. Faraday Trans. 91 (1995): p3255-3261 60 CO OH CO OH P P RT P P RT RT 2 2 ) 73600 exp(10*28.61 ) 59900 exp(10*52.41 ) 18900 exp(10*33.5 8 5 3 ? + ? + = ? ? ? ? (13) An FT WGS kinetic equation is also incorporated into this model of the FT reactor. )( 22 2 2 P HCO COOHwCO K pP pPkr ?= (14) The conversion of CO can be seen in Figure 5 for different promoted iron catalysts that were based on different mechanism (See Chapter 3). Figure 5 CO Conversion in FTS reactor. 0.042l/min, 232C, 20bars H 2 /CO~2 on 40mm I.D Several studies were first conducted on a single tube, varying tube diameter to determine the effects of conversion and temperature rise down the reactor bed. The comparison of a MFEC bed is compared with that of a packed bed. With a constant space velocity and constant flowrate respectively, the temperature & conversion profile between a packed bed and MFEC is shown in Figure 6 & 7 for a tube with constant cooling of 225 o C, and constant particle size of 165um. The packed bed has a temperature rise of 20 o C while 27 Quan-Sheng, L. ?Steady-State and Dynamic Behavior of fixed Bed Catalytic Reactor for Fischer Tropsch Synthesis II. Steady State and Dynamic Simulation Results.? Journal of Nat. Gas Chem. 8 (1999): p238-248 61 that of a MFEC bed is less than 5 o C. With this benefit of MFEC, the tube diameter can be increased and reactor length reduced to achieve the same conversion of CO. This will allow an FTS reactor to be much smaller than conventional reactors. 200 210 220 230 240 250 260 00.511.52 T em p er at u r e ( C ) Length (m) MFEC @ 40mm ID PB @ 15mm ID PB @ 40mm ID PB @ 30mm ID Figure 6 Temperature profile in FTS reactor. Packed bed and MFEC bed contain 165um diameter particle 0 20 40 60 80 100 0 5 10 15 20 % C O C o nver si on Length (m) MFEC -15.3mm ID PB- 15.3mm ID PB @40mm ID MFEC @ 40mm ID Figure 7 CO Conversion profile in FTS reactor for Packed Bed and MFEC. Packed bed and MFEC bed contain 165?m diameter particles. The next set of studies where to compare a conventional packed bed to MFEC reactor where particles sizes were changed to determine the effect of temperature and pressure drop distribution to meet a constant CO conversion. A constant space velocity of 100/hr was enforced. These results are shown in the Table below. 62 Table 6. Conventional Packed Bed Reactors vs. MFEC 70% Conversion CO Packed Bed, 3mm Dp MFEC 165um Tube Diameter, mm 15.3 30 40 15.3 30 40 Length, m 20 16 1 20 13 6 Overall Heat Transfer Coefficient, W/m 2 K 299.96 159.86 121.49 4642.40 2407.20 1880.50 T? / C 31 100 150 2 4 5 Pressure Drop, bar 2 1 - 15 11 6 The following equations, as well as an alpha-temperature relationship shown in chapter 3 allows for simulating and calculating with respect to CO conversion the product distribution in the PFR reactor. A 7 m multi-tubular reactor of 900 tubes with a 2 in I.D is modeled in Aspen to achieve a high pass conversion of 80% for MFEC. With a MFEC bed voidage of 75% and a particle density 1774kg/m 3 , the pressure drop of 7 bar is seen across the reactor. The product distribution of a packed bed and MFEC are shown in figure 8. Figure 8 Product Selectivity Comparisons between Paced Bed and MFEC. A significant improvement is seen with the MFEC in terms of C 9+ . This is because the FTS reactor operates under optimum conditions necessary to yield the highest JP5 selectivity. The packed bed produces more light gases due to the temperature hotspots. The uniformity in temperature is essential to achieving such high C 9 -C 16 selectivity. The advantages can be seen in using MFEC reactor as the reactors yielded higher selectivity to jet fuel and heavier hydrocarbons C 17+ which will be further broken down into smaller chained hydrocarbons via 63 a hydrocracker. FTS reactors provided lower selectivity to light hydrocarbons (C 1-8 ). Un- reacted CO, about 30 mol%, can be recovered from the light gas product stream and recycled back into the reactor after a gas-liquid separation by reacting the gas stream with a lower alkyl-alcohol to form an alkyl-formate. The formate is separated from the fuel gas easily and carbon monoxide is regenerated by decomposition of the formate. 28 Water exiting the FT reactor amounts to about 49 wt.% of all products and contains number of other products that are not modeled; such as organic acids and alcohols. Heavy paraffin conversion process through the use of a hydro-cracking unit is necessary to improve the yield of JP-5. To achieve a high selectivity through this process; the right catalysts and operating conditions have to be selected, so that heavy molecules display much higher reactivity than light components, which will prevent over cracking of materials of C 9 or smaller into light gases. Hydro-cracking catalysts are characterized by the presence of two types of active sites that operate simultaneously; acidic sites, which provides the isomerization/cracking and the metal sites which is responsible for hydrogenation/de- hydrogenation. 29 Common catalysts on amorphous silica-alumina supports are platinum, palladium or bi metallic systems (i.e., Ni/Mo, Ni/W or Co/Mo). Catalysts loaded with a noble metal (particular Pt) show better performances in terms of selectivity for hydroisomerization and product distribution. 30,31 Temperatures of 350-385 o C have shown to favor cracking of heavy hydrocarbons chains in diesel. 32,33 The hydro-cracking unit is operated at a pressure of 28 James, L. ?Recovering carbon monoxide from fuel gas.? U.S Patent -3716619 (1971) 29 Archibadld, R.C ?Catalytic hydro-cracking of aliphatic hydrocarbons.? Industrial & Engineering Chemistry 52 (1960): p745-750 30 Gibson, J.W ?The use of dual function catalyst in isomerization of high molecular weight n-paraffin.? Industrial & Engineering Chemistry 52(1960): p113-116 31 Weitkamp, J. ?Factors influencing the selectivity of hydro-cracking in zeolites.? In: Barthomeuf, D.,, Guidelines for Mastering the Properties of Molecular Sieves. (1990): p343. Plenum Press, New York. 32 Sturtevant, P. et al., ?Fischer-Tropsch Wax Characterization and Upgrading: Final Report.? DOE Report, DE88014638 (1988). 64 50 bars and 1-2 LHSV. 34 The products of the hydro-cracking unit are diesel (80 wt.%), gasoline (15 wt.%) and minimal light gases such as ethane, propane and butane (5 wt.%). 32 It is important to note that diesel and gasoline both contain C 9 -C 16 hydrocarbon chains. For simplicity, the hydrocracker is modeled with a yield reactor, ?RYield,? in AspenPlus TM . Product yields are calculated assuming an 80% conversion of heavy feed (waxes) and the hydrogen consumed in this section is 65% of this heavy feed. The product distributions are shown in figure 9. Figure 9 Products Yield from hydro-cracking. The following table, Table 7, shows the overall yield of JP-5, an improvement is seen when a hydro-cracking unit is included in the process. Table 7 Overall Product Yield of JP-5 Packed Bed wt.% MFEC wt.% Packed bed + Hydro-cracker wt.% MFEC+ Hydrocracker wt.% C 1 -C 8 0.54 0.34 0.6 0.46 C 9 -C 16 0.3 0.34 0.4 0.54 C 17+ 0.16 0.32 - - 33 Fernandes, F. ?Modeling and optimization of Fischer-Tropsch products hydro-cracking.? Fuel Processing Technology 88 (2007): p207-214 34 Walas, Personal Communications (1985) 65 In a further research study, the hydro-cracking unit can be designed to operate with a micro-fiber catalyst, as this should help with the removal as the process is exothermic. The following parameters are suggestions that can be optimized to improve product yield for JP-5 in a hydro-cracking unit; a process configuration layout; the need for a single-stage or two stage reactor or just once-through in order to determine a maximum systematic conversion. In addition to the configuration, the need for and an optimal feed/recycle ratio should be determined as well as the fractional cut point and conversion level in combination with kinetic studies to accurately model this unit. A flowsheet of the downstream process is shown in Figure 10. Figure 10 Process Flow Sheet of FTS - Downstream. 4.2.3 Downstream The fuel and wax is cooled in a series of heat exchangers and a post flash drum is needed to separate light gases and water from the FT reactor products. The stream contains un-reacted CO, CO 2 and C 2 ? C 8 with a heating value of 3.81MMBtu/hr which can be used in the furnace. 2 steam operated distillation columns with 18 stages are used to obtain the 66 96wt% purity of JP-5 at 470 bpd. The waxes are then transported to the hydro-cracking unit. Hydro-treatment of the product is necessary to remove reactive species such as olefins and alcohols that interfere with the hydro-cracking process however; this is beyond the scope of this work. The desire to use one major distillation column to separate and achieve high purity from feeds originating from the hydro-cracking unit and post flash drum will come at a cost of having a high heat duty and a very tall column with 32 stages. 4.3 Weight/Volume BOP Analysis The weight by volume analysis is a unique way at looking at a specific plant design case and this is an important design criterion for a mobile-skid unit. Available space limits how much process equipment one can have. This is why it is important to optimize reactor conditions and have an efficient heat removal system for the FTS reactor to attain a high single pass conversion along the reactor without recycle while maintaining selectivity, resulting in decreased demand on downstream separations. There is however, the need to increase carbon efficiency for larger production units which is achieved through recycling. Piping layout, with isometric drawings and piping and instrumentation diagrams (P&ID) will be the next developmental phase for the skid unit. There are numerous possibilities with regards to FT design, and product slates. However, with regards to only focusing on JP-5, while maintaining minimal volume and weight of the designed plant, three possible case scenarios were developed based on pilot scale-up needs (Case A) and the need for a larger facility capable of producing 500 bpd JP5 (Cases B and C). 67 Figure 11 Fischer Tropsch Possibilities for Mobile Skid Unit. Case A, the once through pass within the FT reactor, is first designed to meet a production target of 2L/12hr day of JP-5 without any recycle. This design is the first stage in a pilot developmental process. A POX, ATR or SMR unit can be used to produce syngas for these cases but the SMR was selected over the others because of overall syngas production output, an oxygen plant is not needed for this design, and the lowest process temperature requirement is used for this process. Case A was enhanced further to meet a target of 500 bpd of JP-5 for a basis of comparisons. This was done assuming that undesirable products, C 1 -C 8 and C 16+, will be burned in the SMR furnace. This is the smallest possible unit that can be designed as there is no need for a hydro-cracking unit. One would not ideally put this design into practice because of the loss of valuable fuel (C 16+ ) associated with this design and its thermal efficiency. Cases B and C, model a more realistic GTL operating plant with the need to recycle lights gases to be burned as fuels in the SMR furnace or reformed to make syngas. Case B models a plant which only focuses on JP-5 as single product produced while Case C 68 allows for the production of JP-5 and light naphtha (C 5 -C 7 ) as its two only products. Both plants have onsite steam producing facilities and a hydro-cracking unit. A simulation model was created to optimize JP5 production for 2 L per 12 hr day with no recycling in a 1.64? I.D tube with a length of 2 m, and this case is illustrated in Figure 12. Figure 12 Process Flow Diagram for Case A (Single Pass). Using a detailed kinetic model in Aspen, the operating temperature was varied, with a constant pressure of 20 bars to find an optimum operating regime to meet the targeted production of 2 L/12-hour day. The major advantage of the use of an MFEC reactor bed in this particular operation is seen with the temperature rise and, selectivity as shown previously in Figure 6. The highest production of JP-5 was achieved at 255 o C with a selectivity of 27wt%. The rate of conversion is much higher within this temperature range. There are disadvantages, however, with performing FT at such high temperatures. The rate of catalyst deactivation is much higher due to carbon deposition from a catalyst phase changes and a possible ?Boudouard? reaction; where carbon monoxide reduces into carbon dioxide and 69 elemental carbon. 35 The temperature rise results in a low ? value and thus, selectivity for C 9 - C 16 hydrocarbons, and an increase in lighter products (C 1 -C 9 ) are seen. As seen in Table 8 and figure 13, JP-5 has a low selectivity of 27wt.%, but the rate of conversion is much higher. Table 8 Case A: Overall Product Yield of JP-5 - HTFT Temp, deg C 235 250 260 270 275 280 Alpha 0.807 0.780 0.762 0.743 0.734 0.725 CO conversion, % 50.2 65.5 73.5 79.6 82.0 83.9 C9-C16 prod., L/12 h day 1.251 1.431 1.478 1.419 1.405 1.323 C17+ prod., L/12 h day 0.442 0.336 0.262 0.220 0.171 0.148 C9-C16 selectivity, % 33.4 29.0 26.6 23.3 22.5 20.7 Figure 13 High Temperature Fischer Tropsch for Case A (Single Pass). To validate our simulation model, it was compared to experimental data that was gathered by Cerametec, an independent R&D group who also have research interests in Fischer-Tropsch catalysis. Figure 14 compares the conversion and selectivity for Case A, where 2L/12 hr day of JP5 is being produced, with that of an iron catalyst at similar experimental conditions. 35 Bohlbro, H. ?An Investigation on the Kinetics of the Conversion of Carbon Monoxide with Water Vapor over Iron Oxide Based Catalysts.? Gjellerup, Copenhagen (1969) 70 0 10 20 30 40 50 210 220 230 % C onv er s ion or S e le c t iv it y Temperature, O C Simulated CO Conv.% Experimental CO Conv.% Simulated JP-5 Sel.% Experimental JP-5 Sel.% Figure 11 Case A (Single Pass) Simulation validation with experimental results. In an effort to model a low temperature Fischer Tropsch (LTFT) to meet the targeted production of 2 L/12 hour day, it was observed that more syngas is needed in the feed stream to meet the production target. Tables 9 and figure 15 illustrate the effects of temperature variation on JP5 production and selectivity for both LTFT reaction regimes. Table 9 Case A: Overall Product Yield of JP-5 - LTFT Temp, deg C 200 205 210 215 220 Alpha 0.872 0.863 0.853 0.844 0.835 CO conversion, % 17.5 20.9 24.9 29.4 34.2 C9-C16 production, L/12 h day 0.506 0.603 0.712 0.821 0.928 C17+ production, L/12 h day 0.275 0.292 0.318 0.342 0.296 C9-C16 selectivity, % 39.1 38.8 38.1 37.4 36.3 With a constant flowrate of 1.4x10 -4 SCMM which is used for the HTFT, a significant amount of JP-5 is not produced, 48% on average, due to low conversion of CO and overall FT reaction rates. However, improved selectivity is maintained at 40 wt.%, which counteracts the reduced reaction rate. 71 Figure 15 Low Temperature Fischer Tropsch for Case A (Single Pass). There are several advantages of running a LTFT for case A: the increased selectivity to JP5 due to operating near optimal ?, reduced soot formation risk compared to high temperature and also a reduced risk of temperature runaway. However, a low single-pass conversion dictates the need for recycle and/or increased feed flowrate. Cases B and C have light gases recycled in order to be burned as fuel in the SMR furnace. There is the addition of a hydro-cracking unit to break down all heavy waxes so as to increase the product yield of JP5. Figures 16 and 17 show block flow diagrams of different case studies for JP-5 production. 72 Figure 16 Process Flow diagram for Case B (Single Pass with Hydro-Cracking Unit). Figure 17 Process flow diagram for Case C (Single Pass with Hydro-Cracking Unit/Lt naphtha). 73 Equipment Sizing In reality, the SMR unit and furnace is a single unit, called heat exchanger reformers, and the heated streams run through piping in the reformer furnace to be heated. 36 A high alloy steel tubes will be used, (25 Cr 35 Ni Nb Ti), 37 known as ?mircoalloys? because of their ability to withstand high temperature and maintain its strength. The shell thickness and tube thickness are 50 mm and 8 mm respectively, which results in a very heavy unit. Assuming a nickel catalyst on zirconia is used for SMR operations which will have WHSV of 1.25 Ibs NGas/hr/Ibcat and catalyst density of 54 Ibs/ft 3 . 38 Detailed kinetics are not included in the model for the SMR unit in Aspen. More details on the mechanical design of steam reformers are available in literature. 39,40 Assuming a liquid hold up of 4 min, 41 the flash drums and distillation columns can be sized. The heat exchangers are sized by calculating the required surface area of heat exchanger from a known heat duty using the overall heat transfer coefficient and log mean temperature difference. m TUFAQ )(?= (14) )]/[( )()( 12 12 log TTlm TT TT meanm ?? ??? =?=? (15) With a shell-and-tube exchanger with multiple-tube passes, a correction factor, F, is applied to compensate for flow direction. 41 Based on a 1.25? square pitch for a 1? tube and two pass 36 Dybkjaer, I. ?Advanced reforming technologies for hydrogen production.? International Journal of Hyrdrocarbon Engineering 3 (1998): p56-58 37 Mohri, T. ?Application of advanced material for catalyst tubes for steam reformers.? Ammonia Plant Safety 33(1993): p86-100 38 Walas, S. ?Chemical Process Equipment Selection and design.? (1988) Butterworth Publishers 39 Dybkjaer, I. ?Tubular reforming and autothermal reforming of natural gas ? an overview of available processes.? Fuel Processing Technology 42 (1995): p85-107 40 Rostrup-Nielen, J.R Catalysis, Science and Technology, (1984) Springer-Verlag, Berlin. 41 Seider, W.D. and Lewin, D.R. ?Process Design Principles.? (2004) 2nd Ed. John Wiley and Sons Inc. New York. 74 heat exchangers of 16 ft, the total number of tubes is estimated from a heat exchanger tube sheet layout count table 19 and with the shell I.D known from the required surface area, each exchanger?s weight and volume can be calculated with an assumed carbon-steel material that has a shell I.D and tube thickness of 0.375? and 0.07? respectively. A more detailed analysis will involve calculating the pressure drop within these exchanger units. All calculations are included in Appendix IV. Table 10, 11 and 12 give a summary of results obtained from Case A, Case B and C. The catalyst bed packing density is 0.98 g/cc for Fe/Cu/K/Al 2 O 3 and MFEC (8 vol.% Cu fiber-15 vol.% catalyst-75 vol.% void) is 0.653 g/cc. Table 10. Products for Case Scenarios Case A Single Pass JP-5 Only Case B Single Pass + HCC JP-5 Only Case C Single Pass + HCC JP-5+Naptha Packed Bed MFEC Packed Bed MFEC Packed Bed MFEC Jet Fuel Fraction (wt.%) 100 100 100 100 60 69 Naphta Fraction (wt %) - - - - 40 31 Case A, no hydro-cracking unit and all off spec. products are burned. Case B, JP-5 only. Case C, JP-5 and light naphtha are products and recycling of lights are burned as fuel. There is the addition of a hydro-cracking unit. Table 11. Weight and Volume of Major Process Equipments for Case A Case A Single Pass JP-5 Only Unit Packed Bed MFEC SMR Weight (tonne) 292 285 Volume (m 3 ) 29 17 FTS Weight (tonne) 495 310 Volume 57 49 75 (m 3 ) Heat Exchanger Weight (tonne) 710 710 Volume (m 3 ) 63 63 Column Weight (tonne) 18.7 16.3 Volume (m 3 ) 13.9 17.4 Table 12. Weight and Volume of Major Process Equipments for Case B and C Case B Single Pass + HCC JP-5 Only Case C Single Pass + HCC JP-5+Lt. Naphtha Unit Packed Bed MFEC Packed Bed MFEC SMR Weight (tonne) 269 256 261 256 Volume (m 3 ) 10 9 10 9 FTS Weight (tonne) 375 128 351 116 Volume (m 3 ) 42 17 21 13 Hydro-Cracking Weight (tonne) 92 90 85 79 Volume (m 3 ) 12 10 9 8.4 Heat Exchanger Weight (tonne) 1247 1021 1247 1021 Volume (m 3 ) 158 130 158 130 Column Weight (tonne) 18.7 16.5 18.7 18.7 Volume (m 3 ) 13.9 11.4 8.6 8.6 Compressors, valves, condensers, pumps, Membrane separator, Hydrotreaters, pipes are not included 76 The effect is drastically seen with the FTS reactor weight and volume. The MFEC reactor tube diameter was scaled up to 50 mm I.D, tube counts for the reactors were varied, and reactor lengths were reduced to 9 m in order to meet a CO conversion of 80% for a single pass. The most significant amount of tubes was seen in the MFEC reactor was for Case A, which had 1500 tubes, compared to MFEC cases B and C with 1100 each. The packed Bed reactors had a tube count of 2050 tubes for each case. At equivalent production rates, MFEC requires much lower flow rates through the steam reformer and the FTS reactor. Smaller heat exchangers are needed and thus smaller reactor sizes. When light gases are recycled to be burned in the SMR, significant improvements can be seen in the reduction of the weight and volume of the overall BOP when compared to a packed bed reactor. Case B offers a reduction in FTS reactor size and an overall 25% reduction is seen when compared to a packed bed. It is important to remove the CO 2 , from the FT tail gas recycle which will improve the thermal efficiency of this process. With light naphtha being produced that can be blended for gasoline production, there is a reduction in light gases that will be burned, and the overall weight and volume of the FTS are significantly reduced. A 30% reduction is seen in Case C. Table 13. General Comparison Cases for 500 bpd of JP-5 for MFEC Cases Case A Single Pass JP-5 Only Case B Single Pass + HCC JP-5 Only Case C Single Pass + HCC JP-5+Lt. Naphtha Packed Bed Single Pass + HCC JP-5+Lt. Naphtha Natural Gas (MMCF/hr) 13.1 6.2 6.2 7.6 Steam (kg/hr) 6200 3400 3400 4200 Weight (tonne) 310 128 116 351 Volume (m3) 49 17 13 21 Lt. Naphtha - - + + 77 Table 14 compares the recycled streams and SMR duty for each case. Case A provides an added advantage of being able to supply heat to the SMR from the off-spec products. Table 14. Composition of Recycled Streams for MFEC Cases Case A Single Pass JP-5 Only Case B Single Pass + HCC JP-5 Only Case C Single Pass + HCC JP-5+Lt. Naphtha Packed Bed Single Pass + HCC JP-5+Lt. Naphtha Natural Gas (MMCF/hr) 0.075 0.3 0.4 0.4 C 1 -C 4 (vol.%) 64.6 34.9 32.7 62 C 5 -C 8 (vol.%) 25.5 22.1 - - C 16+ (vol.%) 99 - - - Heating Value[MMBtu/hr] 7.9 3.81 1.32 4.2 SMR Duty [MMBtu/hr] 49.8 31.3 31.3 35 MFEC bed reactor operates with small particles that greatly improve heat and mass transfer. This allows for a high effectiveness factor for catalyst, maximizing activity and life. This will lead to far high reactor productivity as seen in figure 18. Figure 18 Reactor Productivity. The process by which MFEC is made is by wet lay process, an inexpensive and environmentally friendly paper-based manufacturing process. MFEC FT achieves the 78 economy of scales at much smaller size (500 bpd) than conventional technology (5000 bpd). This advantage allows MFEC-FT to be feasible for GTL applications because of the reduction in its weight and volume. Conclusion A GTL process was modeled and simulated to produce 500 bpd of JP-5. The feasibility of using MFEC for FTS has been verified for a phase I evaluation. Due to the high thermal conductivity of metal fibrous media, MFEC significantly improves the temperature profile in the fixed bed reactor, and allows the FTS take place at lower temperatures. Product selectivity was attributed to maintaining the temperature profile and as a result, JP-5 selectivity was enhanced. The use of a hydrocracker in an FTS process can enhance the selectivity by 30wt.%. It was found that the use of MFEC can significantly reduced the BOP (20 to 30%) and improved the utilization of natural resources (14%) while maintaining the same production capacity. 79 ?Nothing in life is to be feared. It is only to be understood.? Marie Curie Chapter 5 Accomplishments and Future Directions 5.1 Accomplishments The promoted iron catalyst has shown it can have an alpha value of 0.86; high enough to achieve the maximum selectivity towards JP-5 from an FTS reaction. With the use of a micro-fibrous entrapped catalyst, we are able to enhance intra-bed heat transfer by 90%, reducing the temperature rise within the reactor tube, which damages the catalyst. The reduction in temperature rise will ensure a high JP-5 product selectivity in this process. A detailed kinetics has been modeled efficiently with the integration of Matlab and Aspen TM to accurately predict the thermal conductivity of the bed, overall heat transfer coefficient and temperature profile under FTS conditions. This allowed us to accurately model the FTS reactor needed to achieve a conversion of 80%. An overall GTL process was modeled and simulated to produce 2L/12hr day for a small scale pilot plant. The experimental results marched very well with simulated data for 2L/12hr-day for a 1.67 inch, 2 m tube. A larger producing plant, 500 bpd of JP-5, was also simulated. The feasibility of using MFEC for FTS has been verified for a phase I evaluation. Due to the high thermal conductivity of metal fibrous media, MFEC significantly improves 80 the temperature profile in the fixed bed reactor, and allowed the FTS to take place at a low temperature level. Product selectivity was attributed to maintaining the temperature profile; JP-5 selectivity was enhanced. The use of a hydrocracker, allowed us to increase the selectivity of JP-5 by 30wt.%. 5.2 Future Directions There is still a lot of work that can be done on the simulation efforts. First, a pinch analysis could be done to improve heat integration of overall plant system. This could have the potential of reducing the required number of heat exchangers and increase thermal efficiency. Further work needs to be done on modeling hydro-cracking unit with an ideal kinetics. This is an exothermic process where the use of a micro-fibrous material can improve the intra-bed thermal conductivity, thereby allowing one to achieve the same producing capacity with a smaller reactor. An ideal piping layout and P&ID should be constructed with the aid of process control integrations. The next phase of this work involves development of additional process models for the generation of performance metrics, specifically information on conversion, yield, and production cost for economic metrics. This would be analyzed based on recent and future predictions of crude oil and natural gas prices. Although the case study illustrated previously demonstrates real world usage of this methodology, many simplifications have been made, and over time the simplifying assumptions will be further reduced in order to increase the realism and rigor of the framework as an evaluation tool. 81 Bibliography Aasberg-Perterson, K. ?Synthesis gas production for FT synthesis.? Studies in surface science and catalysis. 152 (2004): p258 Aliabad, Z ?Removal of CO 2 and H 2 S using Aqueous Alkanolamine Solution.? World Academy of Science, Engr. 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Science Engineering 41 (1999): p255-315 Walas, Personal Communications (1985) Walas, S. ?Chemical Process Equipment Selection and design.? 2 nd Ed. Butterworth Publishers (1988) Newton Weitkamp, J. ?Factors influencing the selectivity of hydro-cracking in zeolites.? In: Barthomeuf, D.,, Guidelines for Mastering the Properties of Molecular Sieves. (1990): p343. Plenum Press, NY. Yermakova, A. ?Thermodynamica Calculations in the Modeling of Multiphase Processes and Reactors.? Ind. Eng. Chem. Res. 39 (2000): p1453-1472 Appendix I Tapped density is easily measured by placing a powder sample in a volumetric flask and weighing it. The sample is tapped to settle the particles as much as possible. Bulk density is the density of the material without any pore volume. This data is usually available in the literature, handbooks etc. It can be measured by immersing a known weight of sample in a fluid that wets the sample. The bulk density can then be calculated from the increase in volume. Particle density can be calculated from pore volume and bulk density as follows. bulk p particle v ?? 11 += [AP1.] The tap density, bed voidage and particle density are related as follows: () particletap ??? ?= 1 [AP1.2] Bed voidage can be estimated from the tap density as follows: particle tap ? ? ? =?1 [AP1.3] If catalyst or sorbent is loaded on the particles then the total volume of the particles does not change but the weight increases. The formula above can be modified to account for this increased weight. ()L v bulk p particle + += 1 11 ?? [AP1.4] 88 Appendix II Derivation of a Segregated 2-D Model for Heat-up of a Packed Tube with Flowing Gas. The partial differential equations describing the heat balance between gas, catalyst and fibers inside the tube are as follows: nullnull null null nullnull ?T null ? nullnull null 1 r ? nullnull nullr ?T null nullnull nullnullk null ? null T null ?z null nullnull null null nullnull null null ?T null ?z nullx null null1null?nulla null h nullnull nullT null nullT null null t nullx null1null?nulla h nullnull nullT null nullTnull [Ap2.1] null null null x null null1nullnullnullnull null null nullnull nullT null nullnull nullnull null null null null nullnull nullr nullT null nullnull nullnullk null null null T null nullnull null nullx null null1null?nulla null h nullnull nullT null nullT null null [AP2.2] x null null1nullnullnullnull null null nullnull nullT null nullnull nullnull null null null null nullnull nullr nullT null nullnull nullnullk null null null T null nullnull null nullx null null1null?nulla null h nullnull nullT null nullT null null [AP2.3] With boundary conditions: T=T 0 at z=0. And T=T bath at the outside of the tube. x c and x f are the volume fraction of the solid phases occupied by the catalyst and fiber phases (x c + x f =1). The specific surface areas of the catalyst and fibers are designated by a c and a f . The bed voidage is represented by ?. Define:, null null null null null null nullnull null null null nullnull , null null null null null null nullnull null null null nullnull , null null null null null null null null nullnull , null null null null null null null null nullnull , x=z/L and ?=r/R. A heat transfer coefficient is applied at the interface of the tube wall and the gas, catalyst and fiber phases. nullnull null nullT null nullnull nullnull nullnull nullnull null nullnull null null , nullnull null nullT null nullnull nullnull nullnull nullnull null nullnull null null nullnull null nullT null nullnull nullnull nullnull nullnull null nullnull null null [AP2.4] Also, the total heat transferred into the gas, catalyst and fiber phases is equal to that coming through the wall of the tube. 89 null null nullT null nullnull nullnull null nullT null nullnull nullnull null nullT null nullnull nullnull null nullT null nullnull at r=R [AP2.5] A pseudo steady-state assumption for the temperature in the tube wall allows for the effect of the limitation of the thermal conductivity of the tube material to be included in the heat transfer model. 0null null null null nullnull nullr nullT null nullnull null [AP2.6] Given T w =T ? at r=R+W and T w at r=R, estimate nullT null nullnull at r=R and combine this value in Equation 2. W is the wall thickness. nullT null nullnull null null null nullnull null null null null nullnull null? nullT null nullnull nullnull null null nullnull null null null T null nullnull null null null null nullnull nullT null nullnull null null null nullnullnullnullnullnull null null nullT null nullT null nullnull null null nullnullnullnullnullnull null null nullT null nullT null null [AP2.7] nullT null nullnull null null null nullnullnullnullnullnull null null null null nullnull null? nullT null nullnull nullnull null null nullnullnullnullnullnull null null null T null nullnull null null? null nullT null nullT null null [AP2.8] null null null null null null null nullnull null? nullT null nullnull nullnull null null nullnullnullnullnullnull null null null T null nullnull null null? null nullT null nullT null null [AP2.9] nullT null nullnullnullnullnullnullnullnull nullnull null nullnull null nullnull null null null nullnull null nullT null nullnull , nullnull null nullnull null null null null null nullnull null nullT null nullnull and null nullnull null nullnull null nullnull null null null null nullnull nullT null nullnull at ?=1 [AP2.10] -null null nullT null nullnull nullnull null nullT null nullT null nullnullnull null nullT null nullT null nullnullnull null nullT null nullT null null at ?=1, [AP2.11] 0null null null null nullnull null? nullT null nullnull null with T w =T bath at r=R+W [AP2.12] Integrating Equation 12 and applying the boundary conditions results in the following expression for the radial derivative of temperature in the tube wall: nullnull null nullnull null null null null null nullnull null nullnullnullnullnull null null null null [AP2.13] Combining Equations 11 and 13 gives an expression for the wall temperature as a function of the temperatures of the outer wall and the segregated phases: 90 null null null null null nullnullnull nullnull null null nullnull nullnull null null nullnull nullnull null null null null null null nullnullnullnullnull null null null null nullnullnullnull nullnull nullnull nullnull nullnull nullnull null null null null nullnullnullnullnull null null null null at ?=1. [AP2.14] Discretization of the Segregated Energy Balances in Cylindrical Coordinates The radial coordinate is expressed as i =(i-1)* ?and the axial coordinate as x j =(j- 1)* x. A radial average can be obtained by volume averaging from point i-1 to point i+1 as follows: nullnullnull null null null null nullnull nullnullnull nullnull null null null nullnullnullnull null nullnull nullnullnull nullnull null null null nullnull null nullnullnull null null null null null nullnullnull null nullnullnull null nullnullnull null nullnullnull The resulting formula is: nullnullnull null null nullnullnullnullnullnullnull nullnullnull nullnullnullnullnullnullnullnull null nullnullnullnullnullnullnullnull nullnullnull nullnullnullnullnullnull [AP2.15] In a similar manner, a three point approximation for the second order radial derivative can be obtained by volume averaging from point i-1 to point i+1. The resulting formula is as follows: null null null nullnull null? nullT nullnull null= nullnullnullnullnullnullnull nullnullnull nullnullnullnullnullnullnullnull null nullnullnullnullnullnullnullnull nullnullnull nullnullnullnullnullnullnullnull null [AP2.16] A discrete form of Equation 7 can be written using equation 16 for the second order radial derivative of temperature and standard central finite differences for the other terms. Radial averaging using Equation 13 will be used to eliminate the odd/even disconnect that can occur when central finite differences are employed for the first order axial derivative. Replacing the terms in Equations 7, 8 and 9 with central finite differences and averaging the convective flow term over the radial dimension results in the following expressions: null null null,null nullnullnull nullnull null null,null null nullnull null null null null nullnullnullnullnull nullnullnullnullnullnullnullnull null nullnullnull,nullnullnull null nullnull null nullnullnull,nullnullnull null nullnullnullnullnullnullnullnullnullnull null nullnullnull,null nullnullnull nullnull null nullnullnull,null nullnullnull nullnullnullnullnullnullnullnullnullnull null nullnullnull,nullnullnull null nullnull null nullnullnull,nullnullnull null null nullnullnullnullnullnull null 91 null null nullnull null nullnull null null nullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnullnullnull null null,null nullnullnull nullnullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnull nullnull null null nullnull null nullnull null nullnull null nullnullnull,null nullnullnull null2null null null,null nullnullnull nullnull null nullnullnull,null nullnullnull null null nullnull null null nullnullnullnullnullnull null null nullT null null,null nullnullnull null T null null,null nullnullnull null null null null nullnullnullnullnullnull null null nullT null null,null nullnullnull nullT null null,null nullnullnull null [AP2.17] null null null,null nullnullnull nullnull null null,null null nullnull null null null nullnullnullnullnullnull null nullnull null null nullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnullnullnull null null,null nullnullnull nullnullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnull nullnull null null nullnullnullnullnullnull null nullnull null nullnull null nullnullnull,null nullnullnull null2null null null,null nullnullnull nullnull null nullnullnull,null nullnullnull null null nullnull ? null nullT null null,null nullnullnull nullT null null,null nullnullnull null [AP2.18] null null null,null nullnullnull nullnull null null,null null nullnull null null null nullnullnullnullnullnull null nullnull null null nullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnullnullnull null null,null nullnullnull nullnullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnull nullnull null null nullnullnullnullnullnull null nullnull null nullnull null nullnullnull,null nullnullnull null2null null null,null nullnullnull nullnull null nullnullnull,null nullnullnull nullnull ? null nullT null null,null nullnullnull nullT null null,null nullnullnull null [AP2.19] The right side of these equations with all values at the same time are the residuals at point (j,i). The residuals are computed at each time step to monitor the approach to steady- state. A line implicit solution in the axial direction will be used to integrate the discrete equations to the steady-state solution. This results in a block tri-diagonal system of linear equations: nullnull null null nullnullLnullnull null null null nullL null nullnull null nullnull null nullnullnull,null nullnullnull nullnull null nullnull null nullnull null null null null R null nullnull null null null L null nullnull null nullnull nullnullnullnullnullnullnull null null null nullnull null null null null null nullnull null null,null nullnullnull null nullnullnullnullnullnull null null null null ? null null null null,null nullnullnull null nullnullnullnullnullnull null null null null ? null null null null,null nullnullnull nullnull null null nullnullLnullnull null null null nullL null nullnull null nullnull null nullnullnull,null nullnullnull null null null null,null null nullnull null null null nullnullLnullnull nullnullnullnullnullnullnullnullnull null nullnullnull,nullnullnull null nullnull null nullnullnull,nullnullnull null nullnullnullnullnullnullnullnullnullnull null nullnullnull,nullnullnull null nullnull null nullnullnull,nullnullnull null nullnull nullnullnullnullnullnull null null null nullR null nullnull null null nullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnull null [AP2.20] 92 nullnull null null nullnullnullnullnullL null nullnull null nullnull null nullnullnull,null nullnullnull null? null null null null,null nullnullnull nullnull null nullnull null nullnull null nullnullnullnullnull null null R null nullnull null null null L null nullnull null nullnull? null nullnull null null,null nullnullnull nullnullnull null null nullnullnullnullnullL null nullnull null nullnull null nullnullnull,null nullnullnull null null null null,null null nullnull null null null nullnullnullnullnullR null nullnull null null nullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnullnullnull null null,nullnullnull null nullnullnullnullnullnull null [AP2.21] nullnull null null null1null?nullL null ?x null nullnull null nullnullnull,null nullnullnull null? null null null null,null nullnullnull nullnull 1 ?t null 2null null null1nullnullnull null 1 R null ?? null null 1 L null ?x null nullnull? null nullnull null null,null nullnullnull nullnullnull null null null1null?nullL null ?x null nullnull null nullnullnull,null nullnullnull null null null null,null null ?t null null null null1null?nullR null ?? null null null2null null3nullnull null null,nullnullnull null nullnull2null null1nullnull null null,nullnullnull null 2nullnullnull1null null [AP2.22] The boundary condition at the inlet for these energy balances is a fixed temperature. The zero axial derivative at the exit of the tube is accounted for using a one-sided three point derivative. null null,null nullnullnull nullnull nullnullnullnullnull and null null nullnullnull,null null null4null null null,null nullnullnull null3null null nullnullnull,null nullnullnull 2?null null null0, k=gas, catalyst and fiber. This set of 3x(M+1) equations was solved for different radial coordinates sweeping from the tube wall to one point away from the centerline. The solutions were updated as soon as new values were computed. At the centerline the radial derivative is zero, so the values at the second radial point were used for the centerline values. A three point one-sided derivative that maintains second order accuracy is as follows: nullnull nullnull null null nullnullnull nullnullnull nullnullnull nullnullnull null nullnullnull [AP2.23] Inserting Equation 23 into Equation 10 for either the gas, catalyst or fiber phase gives a formula for updating the temperature of that phase at the wall: nullnull null nullnull null null,null nullnullnull nullnull null null null nullnull null null null null,nullnullnull null nullnullnull null null,nullnullnull null nullnullnull null null,null nullnullnull nullnullnull 93 Rearranging this equation results in the following expression: null null null,null nullnullnull null null null null null null null null nullnull nullnullnullnull nullnullnull null null,nullnullnull null nullnull null null,nullnullnull null null nullnullnull null null null nullnull nullnullnullnull [AP2.24] The k subscript represents the gas, catalyst or fiber phase. 94 Appendix III AP Figure 1 Phase Diagram for Iron Magnetite i The following figure is used to illustrate the trends and possible phase change with respect to temperature for Iron. At a LTFT, the risk of phase change is minimal. The risk increases when running a HTFT. i Muan A. ?Phase Equilibria at high temperatures in Oxide Systems involving Changes in Oxidation States.? American Journal of Science 256 (1958): p171-207 95 Appendix IV Steam Reactor and Furnace The Steam Reforming Reactor was modeled as a fired heat exchanger. Where the numbers of tubes and cross sectional area were calculated using the following equations listed below. The following results are for Case B as an example Ap Table 1. Summary of Size SIZING: Packed Bed MFEC WHSV [lbs NGAS/hr/lb Cat] 1.25 1.25 Cat Density [lbs/ft^3] 54 54 Flow Rate NGAS [lb/hr] 7200 6821 Amount Catalyst [lbs] 5760 5457 Volume Catalyst [ft^3] 107 101 Bed Voidage 0.45 0.45 Steam Reactor Volume [ft^3] 194 184 Flowrate [Ft^3/hr] 5379000 49724 Flowrate [Ft^3/s] 1494 1494 Volume, [Cubic ft] 75 69 Diameter, [Ft] 8 8 Height, Ft 4 4 Number of Tubes: 9864 9864 Tube Diameter [ft] 0.167 0.167 Surface Area of Stm Reactor [ft ^2] 65 65 Shell Wall Thickness, [ft] 0.5 0.5 Tube thickness, [ft] 0.0262 0.0262 Total Weight [Tonne] 269 256 Amount of Catalyst = WHSV asFlowrateNg (AP4.1) 96 BedVoidage talystvolumeofCa ?1 Steam Reactor Volume = (AP4.2) Area of Reactor = actorLengthofre orVolume (AP4.3) Steamreact ii 3.0*TubeArea torAreaofreac N = (AP4.4) The length of the tubes was calculated fr tubes om the reformer volume. By specifying ter, the cross sectional area was calculated. The Steam Reformer consists FTS Reactor the tube diame of a fired furnace which is ignored. Ap Table 2. Summary of Size Packed MFEC Bed Number of Tubes: 2050 1100 Tube Diameter [ft] 0.167 0.167 Tube Length [ft] 42 30 Tube Volume [ft^3] 0.86 0.44 Total Tube Volume 1 3 (w/o catalyst voidage) [ft^3] 758 93 Catalyst Voidage 0.38 0.75 Reactor Volume [ft^3] 949 110 Reactor Volume, [m^3] 27 3 Cat Density [lbs/ft^3] 943 1500 Amount Catalyst [lbs] 55904 9198 Wall Thickness, [ft] 0.5 0.5 Density of Steel, [Ib/ft 553 553 3] Total Volume,[ft3] 1496 597 Total Wight, [Tonne] 375 128 The following formulas lis reactor and its size. ted below are the calculations that were made in order to figure out the cost of the FTS (AP4.5) LrTv 2 ?= ii Seider, W.D. Seader, J.D. and Lewin, D.R.. Process Design Principles. John Wiley and Sons Inc. NY. 2nd Ed. 2004. 97 Tube Volume tubestoutcatalysw NTvTv * / = Total tube Volume w/out Catalyst (AP4.6) ? v R / = lengthdiametertubesA TTNR ** toutcatalysw Tv ?= (AP4.7) Distillation Towers In order, evaluate the distillation column size, certain assumptions were made. These assumptions include no foaming; the sieve trays area ratio is greater than 0.1. The is 24 in. spacing between the trays. A 4 minute liquid hold up was assumed in the reflux accumu r drums. Ap Table 3. Summar Dimension w Reflux A final assumption was that there lato y of Size s of Distillation To er ccumulator Drum Fraction, f 0.8 Flow rate, Ib/Min 16118 Foaming Factor, F F Li 1 quid Hold Up 4 Hole Area factor, F HA 1 Volumetric Flow r 6ate cft 4472 Surface tens Vessel Volume cft 107453. ion, ? H2O (dyne/cm) 72.7 Surface ten Length (Ft) 85.508 sion, ? HC (dyne/cm) 22.7 ? (dyne/cm) 17.26 Diameter (ft) 40 Surface Tension factor, F ST 0.97 Ratio 2.13 Den 3 sity, ? G (g/cm ) 0.0051 Density, ? L (g/cm 3 ) 0.683 Gas, mass velocity, G (g/s) 100044 Liquid, mass velocity, L (g/s) 84629 Flow rat 0. 3 io, F LG 07 Flooding Correlation for sieve, C SB (ft/s) 0.37 Capacity, parameter, C (ft/s) 0.359 Vapor flooding velocity, U f (ft/s) 4.126 D T (ft) 7 H (ft) 36 E, fractional Weld efficiency 0.850 d d P PSE DP t 2.12 ? = This is the wall thickness [inches] (AP4.8) 98 ( )[ ] 2 0 ))(ln(*0015655.0)(*91615.0*91615.060608.0exp Od PPLnP ++= Po is the operating pressure. In (AP4.9) order to calculate the diameter of the Column, the following equation was used; 2 1 )1( 4 ? ? ? ? ? ? ? ? ? ? ? ? ?? = G T d f T A A FU G D ? (AP4.10) where 2 1 ? ? ? ? G f ? ? ? ? ? ? = Gl CU ?? (AP4.11) r is for tow We assumed The paramete ers with perforated sieve plates. 1.0)( = T A d A because the flow ratio, parameter was less than 0.1 Where the weight is related to the thicknes LG F s of the shell, and Height and diameter of the column ))(8.0)(( ? SS tDLtDW ++?= (AP4.12) Ap. Tab stillation Cle 4. Results of Di olumn Size PB MFEC C 7 ft di eter 5 ft di eter OLUMN COST am am Wall thickness, ts [ft] 0. 0. 3 0833 083 Column Diameter, Di [ft] 7 5 Column Height, L [ft] 36 36 p (CS) [lb/ft^3] 490 490 Column weight, W [lb] 37785 11733 99 Flash It is assumed that the material of construction will be carbon steel. Ap. Table 5. Results of Distillation Column Comparison SIZING: Umax, Ft/Sec 3.63 3.51 K, Ft/sec 0.35 0.35 Pl, liquid Density, Ib/Ft^3 35.83 46.75 Pv, Vapor Density, Ib/FT^3 0.33 0.46 Vapor Rate Ft^3/hr 134400 1183000 Vapor Flowrate Ft^3/s 37.33 328 Diameter, Ft 3.61 10.916 Height, Ft, Hl 5.46 0.210 Height Hv 4 4 4 Total Height or Length [Ft] 9.45 4.21 t, Liquid Hold up, [mins] 4 4 Liquid Flow rate, [Ft^3/s] 3.87 0.510 Thickness, [in] 1.440 1.87 Pd, internal Design [Psig] 813.5 358 Di, inside diameter, [Inches] 43.42 130 S, Max allowable Stress of Shell Material, [psi] 15000 15000 E, fractional Weld efficiency 0.85 0.85 Po, Operating Pressure, [psig] 720 300 Temperature inside the flash, [F] 100 150 Max Vapor Velocity 2 1 *max ? ? ? ? ? ? ? ? ? = V vl P PP KU K = 0.35 w/Demister (AP4.13) 2 1 max 4 min ? ? ? ? ? ? ? ? ? = U Q D v (AP4.14) Overall , Height = H l + H v. Assumed , Liquid Holdup, is about 4 mins l t ? ? ? ? ? ? ? = 2 * **4 * D Qt KH ll l (AP4.15) 100 Ap. Table 6. Summary weight of Flash Columns The Flash Separation Flash H20 Flash Total volume is [Ft^3} 97 394 Thickness [Ft] 0.12 0.156 Head Volume [ft^3] 10 93 Shell [Ft^3] 12 22 Total Vol [Ft^3] 33 209 den of SS [Ib/cft] 505 505 W [Ib] 16903 105921 Heat Exchangers The following are sample calculations use. The Total spreadsheet cannot be shown Ap. Table 7. Summary weight of Flash Columns B22 B23 Q[Btu/hr] 5,701,431 10,855,268 T_process_1[F] 90 100 T_process_2[F] 950 950 T_medium_1[F] 1384 1327 T_medium_2[F] 1327 1290 DT1 434 377 DT2 1237 1190 DT2-DT1 803 813 ln(DT2/DT1) 1 1 T_mean 767 707 R 0.1 0.04 S/P 15 23 Ft 1 1 U[Btu/hr-sqft-F] 150 150 A[ft2] 50 79 Shell Radius[ft] 4 5 total no of tubes 496 804 Shell ID thickness[ft] 0.03125 0.03125 tube tickness[ft] 0.00583 0.00583 p (CS) [lb/ft^3] 490 490 Tube Weight[Ibs] 24281 39358 Shell Weight[Ibs] 98908 157532 Weight[Ibs] 123189 196890 Weight[tonne] 56 89 Volume[ft3] 251 402 Volume[m3] 7 11 101 Hydro-cracking Ap. Table 8. Summary Hydro-cracking Unit HYDROCRACKING Hydrocracking yield(%) 80 WAX input to the process (kg/h) 750 residue unreacted (kg/h) (recycle) 150.00 total wax input (kg/h) 900.00 Hydrogen that reacts (kg/h) 4.5 Hydrogen excess (kg/h) 1.35 total H2 input reactor (kg/h) 65% of total wax input, weight 5.85 TOTAL INPUT (kg/h) 905.85 with a yield of 5% (weight) for gas, 15% gasoline and 80% diesel diesel produced (kg/h) 603.60 nafta produced (kg/h) 113.18 gas produced (kg/h) only saturated gases 37.73 TOTAL OUTPUT (kg/h) 755.85 Ap. Table 9. Summary Hydro-cracking Unit Variables Used in Aspen distribution diesel (saturated hydrocarbon) PM (kg/kmol) weight fraction kg/h (reactor output) Yield (*) C12 170 0.1627 98.23 0.1084 C13 184 0.1587 95.76 0.1057 C14 198 0.1508 91.02 0.1005 C15 212 0.1426 86.09 0.0950 C16 226 0.1398 84.39 0.0932 C17 240 0.1298 78.32 0.0865 C18 254 0.1156 69.78 0.0770 tot 1.000 603.60 gas kg/h produced Yield CH4 9.43 0.0104 C2H6 9.43 0.0104 C3H8 9.43 0.0104 C4H10 9.43 0.0104 tot 37.73 102 distribution gasoline weight fraction kg/h produced Yield C5 0.014 1.61 0.0018 5 C6 0.028 3.21 0.0035 6 C7 0.065 7.40 0.0082 7 C8 0.146 16.48 0.0182 8 C9 0.237 26.79 0.0296 9 C10 0.253 28.64 0.0316 10 C11 0.257 29.04 0.0321 11 tot 1.000 113.18 tot distribution atoms of C weight fraction Yield 19 0.130312751 0.0216 20 0.103849238 0.0172 21 0.086206897 0.0143 22 0.071371291 0.0118 23 0.059743384 0.0099 24 0.052927025 0.0088 25 0.04691259 0.0078 26 0.042502005 0.0070 27 0.037690457 0.0062 28 0.034482759 0.0057 29 0.030473136 0.0050 30 0.028468324 0.0047 31 0.026463512 0.0044 32 0.023255814 0.0039 33 0.021251002 0.0035 34 0.020048115 0.0033 35 0.018043304 0.0030 36 0.016439455 0.0027 37 0.015236568 0.0025 38 0.014033681 0.0023 39 0.013231756 0.0022 40 0.011627907 0.0019 41 0.010825982 0.0018 42 0.009222133 0.0015 43 0.008821171 0.0015 44 0.008420209 0.0014 45 0.008019246 0.0013 46 0.007217322 0.0012 47 0.006014435 0.0010 48 0.005613472 0.0009 49 0.004811548 0.0008 50 0.004410585 0.0007 51 0.004009623 0.0007 52 0.003608661 0.0006 53 0.002806736 0.0005 103 54 0.002405774 0.0004 55 0.002004812 0.0003 56 0.002004812 0.0003 57 0.002004812 0.0003 58 0.001202887 0.0002 59 0.000801925 0.0001 60 0.000801925 0.0001 61 0.000400962 0.0001 104 Appendix V Fe/Cu/K/La- Al 2 O 3 105 Fe/ Al 2 O 3 106 107 Fe/SiO 2 /Al 2 O 3 Appendix?VI? FTSFEED STM NGAS2 SMRFEED H20 NGAS1 CO2 SMRPROD 11 AIRFURN NGFURN 9 12 SYNGAS1 CW1 STEAM CLNH2O SYNGAS2 18 FTS2 STMMIX SMR SMRFURN C1F1 C1F2 H2OFLASH B2 CRFTS CO2IN B5 SYNGAS3 RCO2CH4 15 FURNFUEL LFUEL 6 2RCO2 AB5 B11 2RH2O GAS B20 B21 26 AB51 3RCO2 B18 24 B19 RCO SMRFUEL B22 B23 5 B24 1 B25 B27 4 B28 B29 14 B31 B32 32 ? Aspen Flowsheet Upstream 108? ? Temperature (C) Pr es s ure (ba r) Mass Flow Rate (kg/hr) V olume Flow R ate (bb l/da y) V apo r Fr acti on Warnings LTGFLASH CRJP5 H2O DC1 JP5 WAX 17 CRACKER B6 WAX2 2CRACKER CRACKPDT PUREH2 B9 2 10 B14 22 DC2 DC2BTMS 2JP5 STEAM2 28 B15 29 30 B17 31 PDT B12 D B33 B34 36 37 B35 B36 39 40 B37 41 42 43 NAPTHA Aspen Flowsheet Downstream 109? ?