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The feasibility of using micro-fibrous entrapped catalysts in Fischer-Tropsch Synthesis for selective JP-5 production as well as overall balance of plant reductions (BOP) will be addressed in this article. 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 210oC 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 mobilized GTL plant design for offshore operations.
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. Other reasons, especially those pertaining to the military, include 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. 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 are smaller or rapid depletion is indicated, as constructing a gas pipeline would present a large financial risk.,,,,,
The objective of this research 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 a metal micro-fibrous entrapped catalyst (MFEC) in a fixed bed reactor with the goal of producing JP-5. 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 which would reduce the BOP requirements for post processing operations. 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.
Due to the exothermic nature of FTS, a fundamental need for applied innovative research into catalytic materials as well as structural architectures and process designs that enable FTS to be transformed into a modular and inexpensive technology is needed. This can be achieved through a MFEC, a catalyst structure which possesses high inherent heat transfer rates offering the opportunity for a better control on intra-bed hot spots and product selectivity. MFEC, a catalyst structure with sintered metal micro-fibers, were developed and patented by Auburn University.,, MFEC have unique properties such as high void volume, enhanced mass transfer, ultra-high contacting-efficiency and a high thermal conductivity when metal fibers are used. These attractive properties are very promising and can be used in different applications: sorbents for carbon monoxide and hydrogen sulfide removal,, electrocaltalyst, filtration, catalytic reactions, and in this particular article, FTS. The use of an MFEC reactor in FTS will not be subjected to drawbacks that conventional fixed bed reactors face, (e.g. hot spots and high gas compression cost).
Integrated Simulated Designs
Aspen PlusTM 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: O2, N2, CO, CO2, H2, and paraffins from CH4-C30H62 with a minimum number of other heavy hydrocarbons introduced into Aspen because of its lack of property definitions for higher chain alkanes greater than C31+. 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 (C5-C11), JP-5 (C9-C16) and waxes (C20+). 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 1. 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.
Figure 1 General Flow Scheme of an FTS Process
The manner in which syngas is produced, steam methane reforming (SMR), partial oxidation (POX) or auto-thermal reforming, 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. There have been extensive research and comparisons on syngas technologies., 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. Table 1 compares syngas generating facilities to one another, highlighting SMR as a favorable choice of design for this operation.
Table 1. Comparison of Syngas generation technologies.
Operating Temperature, oC
Thermal Efficiency, %
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 the flexibility to produce more hydrogen. Another attractive reason for choosing SMR as the favorite syngas producing technology is an oxygen producing 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 scales (>10,000bpd) GTL facilities because of the economics but this technology is sufficient for a 500 bpd plant.17
Synthesis gas for low temperature Fischer Tropsch (LTFT) operations has an optimum H2/CO ratio slightly greater than 2 with low CO2 concentration. However, hydrogen starved syngas increases chain growth probability, and thus selectivity hence, a target ratio of H2/CO of 1.9 from the SMR unit will be set. A process flow diagram of the upstream process is shown in Figure 2.
Figure 2 Process Flow Diagram of FTS - Upstream.
The key reactions modeled in the SMR unit are;
CH4 + H2O â†’ CO + 3H2O (2)
CH4 + CO2 â†’ 2CO + 2H2O (3)
CO + H2O â†” CO2 + H2 (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% N2, 96% CH4, 1.5% C2H6 and 0.5% C3+ (by mole fraction).16 The furnace is modeled as two blocks in Aspen, with one being the reactor (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 900oC. 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 500oC and natural gas (stream NGAS1) is preheated to 500oC at 20 bars using the heat from the furnace. These streams heat up to just slightly above 500oC, the minimum temperature suggested by IEA and react through the 900oC 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 CO2 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 CO2, causes a water gas shift reaction to occur, reducing the ratio of H2/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 H2/CO ratio of 1.9 of the SMR unit (results displayed in Table 2). 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 CO2 recycle favors the desired ratio of H2/CO but does not increase the thermal efficiency. This is can be attributed to the effect of the recycled CO2. 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.
Table 2. Sensitivity Analysis of steam to methane ratio and N.G temperature on GTL Performance.
Thermal Efficiency %
N.G temperature oC
Thermal Efficiency %
After the syngas exits the steam reformer, it is cooled down to 70oC 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 CO2 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, or through the use of N-methyl-diethanolamine (MDEA). MDEA is preferred as it is less energy intensive and operates at a favorable temperature of 57oC. 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.
The FTS reaction is highly exothermic and has an adiabatic temperature rise of up to 1750oC. 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 but there are problems associated with these as well, such as low volumetric reactivity and flow distribution. MFEC reactors combine the advantages of a fixed and a slurry reactor without their corresponding disadvantages assuming operating equipments are properly maintained. Table 3 summarizes the advantages and disadvantages of each of these reactors and highlights the benefits of a MFEC reactor.
Table 3. MFEC Reactors versus Fixed and Slurry Reactors.
Fixed bed reactor
MFEC Fixed bed reactor
Bubble Column reactor
Catalyst content in reactor
Gas-liquid mass transfer
Need for liquid-solid separation
(+) 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. The 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:
The wall heat transfer coefficient with respect to fluid and solid respectively:
The effective radial thermal conductivity is calculated and used in solving an overall heat transfer coefficient.
The Overall Heat transfer coefficient:
The relationship for concentration and temperature dependence in a 1D heterogeneous model are listed below:
A recent kinetic study by Liu allows for the use of intrinsic kinetics for FT synthesis over the design catalyst of choice, Fe-Cu-K. The lumped reaction kinetic by Liu is listed below.
The effectiveness factor on the Fe-Cu-K catalyst has been reported by Wen Jie
An FT WGS kinetic equation is also incorporated into this model of the FT reactor.
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, the temperature profile between a packed bed and MFEC is shown in Figure 3 for a tube with constant cooling of 225oC, and constant particle size of 165um. The packed bed has a temperature rise of 20oC while that of a MFEC bed is less than 5oC. 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.
Figure 3 Temperature profile in FTS reactor.
Constant space velocity ~100/hr Packed bed and MFEC bed contain 165Âµm diameter particle
Figure 4 CO Conversion profile in FTS reactor for Packed Bed and MFEC.
Constant space velocity ~100/hr. Packed bed and MFEC bed contain 165Âµm 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.
Table 4. Conventional Packed Bed Reactors vs. MFEC
70% Conversion CO
Packed Bed, 3mm
Tube Diameter, mm
Overall Heat Transfer Coefficient, W/m2K
Pressure Drop, bar
The following equations, as well as an alpha-temperature relationship 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 model in Aspen to achieve a high pass conversion of 80%. With a MFEC bed voidage of 75% and a particle density 1774kg/cum, the pressure drop of 7 bar is seen across the reactor. The product distribution of a packed bed and MFEC are shown in Figure 5.
Figure 5 Product Selectivity Comparisons.
A significant improvement is seen with the MFEC in terms of C9+. 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 rise within the reactor. The uniformity in temperature is essential to achieving such high C9-C16 selectivity. The advantages can be seen in using MFEC reactor as the reactors yielded higher selectivity to jet fuel and heavier hydrocarbons C17+ which will be further broken down into smaller chained hydrocarbons via a hydrocracker. FTS reactors provided lower selectivity to light hydrocarbons (C1-8).
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 C9 or smaller into light gases. Extensive research have been done on hydro-cracking units.,,,, The hydro-cracking unit simulated will operate at a temperature and pressure of 350oC 50 bars respectively and 1-2 LHSV. 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.%). It is important to note that diesel and gasoline both contain C9-C16 hydrocarbon chains. For simplicity, the hydrocracker is modeled with a yield reactor, "RYield," in AspenPlusTM. 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 6.
Figure 6 Products Yield from hydro-cracking.
The following table, Table 5, shows the overall yield of JP-5, an improvement is seen when a hydro-cracking unit is included in the process. A process flow diagram of the downstream process is shown in Figure 7.
Table 5 Overall Product Yield of JP-5
Packed bed + Hydro-cracker
Figure 7 Process Flow Sheet of FTS - 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, CO2 and C2 - C8 depending on process configuration, which will be sent to the furnace. 2 steam operated distillation columns are used to obtain the 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.
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. 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). A general FT scheme for a mobile skid unit is shown in Figure 8.
Figure 8 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. 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, C1-C8 and C16+, 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 (C16+) 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 allows for the production of JP-5 and light naphtha (C5-C7) 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 9.
Figure 9 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 255oC 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 elemental carbon.38 The temperature rise results in a low Î± value and thus, selectivity for C9-C16 hydrocarbons, and an increase in lighter products (C1-C9) 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 6 Case A: Overall Product Yield of JP-5 - HTFT
Temp, deg C
CO conversion, %
C9-C16 prod., L/12 h day
C17+ prod., L/12 h day
C9-C16 selectivity, %
Figure 10 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.
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 7 Case A: Overall Product Yield of JP-5 - LTFT
Temp, deg C
CO conversion, %
C9-C16 production, L/12 h day
C17+ production, L/12 h day
C9-C16 selectivity, %
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.
Figure 12 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 Î±, 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.
Figure 14 Process Flow diagram for Case B (Single Pass with Hydro-Cracking Unit).
Figure 15 Process flow diagram for Case C (Single Pass with Hydro-Cracking Unit/Lt naphtha).
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.38 A high alloy steel tubes will be used, (25 Cr 35 Ni Nb Ti),39 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/ft3.40 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.41-42 Assuming a liquid hold up of 4 min,43 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.
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 heat exchangers of 16 ft, the total number of tubes is estimated from a heat exchanger tube sheet layout count table19 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/Al2O3 and MFEC (8 vol.% Cu fiber-15 vol.% catalyst-75 vol.% void) is 0.653 g/cc.
Table 8. Products for Case Scenarios
Single Pass+ HCC
Single Pass+ HCC
Jet Fuel Fraction (wt.%)
Table 9. Weight and Volume of Major Process Equipments for Case A
Table 10. Weight and Volume of Major Process Equipments for Case B and C
Single Pass + HCC
Single Pass + HCC
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 CO2, 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 12 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 11. General Comparison for 500 bpd of JP-5 - MFEC Cases only
Single Pass + HCC
Single Pass + HCC
Natural Gas (MMCF/hr)
Lt. Naphtha [MMBtu/hr]
Table 12. General Comparison for 500 bpd of JP-5 - MFEC Cases only
Table 14. Composition of Recycled Streams for MFEC Cases Only
Single Pass + HCC
Single Pass + HCC
Natural Gas (MMCF/hr)
SMR Duty [MMBtu/hr]
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 16 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 realizes economies of scale 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.
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.