Catalytic Partial Oxidation Pilot Plant Study - Industrial & Engineering

Nov 24, 2009 - The experimental work was performed at Praxair's laboratory, process design and economic estimates were performed by Foster Wheeler, an...
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Ind. Eng. Chem. Res. 2010, 49, 94–103

Catalytic Partial Oxidation Pilot Plant Study Vasilis Papavassiliou,* Perry Pacouloute, KT Wu, and Raymont Drnevich Praxair Technology Center, 175 East Park DriVe, Tonawanda, New York 14150

Dionisios Vlachos UniVersity of Delaware, Newark, Delaware

John Hemmings† and Leo Bonnel‡ Foster Wheeler Corporation, 585 North Dairy Ashford Street, Houston, Texas 77079

A new high-pressure catalytic partial oxidation reactor was developed for high-pressure production of synthesis gas from natural gas and pure oxygen, and the integration of the new reactor with the gas to liquids process was investigated. The experimental work was performed at Praxair’s laboratory, process design and economic estimates were performed by Foster Wheeler, and theoretical simulations were preformed by University of Delaware. The reactor system was based on a millisecond contact time mixer capable of delivering a fueloxygen mixture to the catalyst without onset of homogeneous reactions. Significant differences were observed between high- and low-pressure operation. With proper mixer and catalyst bed design the system was able to operate stably at pressures up to 18 bar with performance comparable to low-pressure operation. Process economic analysis predicted a small economic advantage for a catalytic partial oxidation reactor compared to conventional technology. 1. Introduction Synthesis gas is the raw material used for the production of key chemicals such as hydrogen, ammonia, acetic acid, and oxoalchohols. It is also used in the production of liquid fuels from natural gas via Fischer-Tropsch (F-T) process (gas-to-liquids or GTL). The three major commercial processes for synthesis gas production in operation today are (a) steam methane reforming (SMR), (b) noncatalytic partial oxidation (POX), and (c) autothermal reforming (ATR). SMR is the most widely used technology for on-purpose H2 production.1 Approximately 10% of the natural gas used in steam reforming is burned to produce sufficient heat to reform the remaining fuel with steam over a nickel catalyst. Steam reforming produces a synthesis gas containing 3:1 H2/CO ratio. POX, as typified by the GE gasification and shell gasification processes,1 introduces natural gas and oxygen into the reactor by means of a specially designed burner. No catalyst is present in the reactor and the reaction is purely homogeneous. Oxygen is consumed near the reactor entrance to form steam, CO2, and heat which, in the main reactor, are used to reform the remaining methane. The residence time in the reactor is about 1-4 s. The raw syngas that is produced by POX contains a 1.8:1 H2/CO ratio.2 ATR, typically a Haldor Topsoe design, combines partial oxidation and steam reforming to produce synthesis gas and is the preferred technology today for gas-to-liquids applications.2 Sulfur free natural gas is mixed in a burner with oxygen, steam, and optionally a recycle stream containing CO2. The O2 is consumed in this first reaction zone. The hot gas from the burner is fed to a nickel-based reforming catalyst bed below the burner where the final equilibration takes place. The raw syngas contains a H2/CO ratio close to 2.3:1 without CO2 recycle or close to 2:1 * To whom correspondence should be addressed. E-mail: [email protected]. † Current address: 22806 Chaus Court, Katy, TX 77494. ‡ Current address: SNC Lavalin, 9009 West Loop South, Houston, TX 77096.

with CO2 recycle. A H2/CO ratio of 2 is preferred for Fischer-Tropsch synthesis. In the recent past there has been considerable work on high temperature, short reaction time, catalytic partial oxidation processes. The primary difference of this process compared to SMR, ATR, and POX is that the heat required by the endothermic reaction is not generated in separate parts of the reactor but instead heat generation and reaction occur simultaneously on the catalyst. Schmidt pioneered this work at the University of Minnesota using highly reactive monolith reactors coated with Pt or Rh and has reported3,4 greater than 90% H2 and CO selectivities at high methane conversion at atmospheric pressures. The reaction times were on the order of a millisecond, no steam was used, and the syngas produced had a 2:1 H2/CO ratio. Suitable catalysts for syngas production by catalytic partial oxidation of methane are generally noble metals supported on ceramic supports.3,4 Platinum, rhodium, ruthenium, and nickel have good catalytic activity and selectivity for syngas.3,4 Catalysts can be prepared in the form of pellets and monolithic structures with either honeycomb or foam structure. Ceramic supports commonly used are silica, alumina, zirconia, and combinations of those. Metallic monoliths wash-coated with ceramic porous films and impregnated with the precious metal catalyst have also been used. Generally the metal represents 2-10% of the catalyst weight. Rhodium on alumina was found to be the most active and selective metal-support combination, and the washcoated monoliths performed better than the wet impregnated ones.5 Several commercially available catalysts (pellets) were also studied in our laboratory for methane catalytic oxidation. A platinum catalyst supported on γ-Al2O3 pellets (0.125 cm diameter) supplied by UOP (UOP LLC, Des Plaines, IL) was shown to have superior performance over other commercial pellet catalysts6 and was also used in this study. Catalytic partial oxidation is a potential fit for GTL applications primarily for three reasons: (a) elimination of steam to the reactor feed which lowers the oxygen requirements, (b)

10.1021/ie900872r  2010 American Chemical Society Published on Web 11/24/2009

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elimination of carbon dioxide recycle which lowers oxygen consumption, and (c) compact reactor size which reduces capital cost. However, studies carried out in small laboratory-scale units did not address the safety aspects of mixing O2 and hydrocarbons under high temperature and pressure conditions. Reyes et al.7 outline the main challenges for industrial scale catalytic partial oxidation systems as follows. (1) Heat losses from small laboratory systems are significant and affect conversions and selectivities, thus, extrapolations to larger systems are not straightforward. (2) Hydrocarbon/oxygen mixing must be achieved without the onset of homogeneous reactions. (3) Operation at high pressure requires higher temperatures to maintain yield To reduce heat losses and operate at high pressures an internally insulated stainless steel vessel is required. To overcome the onset of homogeneous reactions a compact mixing system located as close to the catalyst as possible is necessary. This is accomplished in this study with a new reactor system named the catalytic hot oxygen reactor (CHOR). CHOR is based on a thermal nozzle8 that produces a very high temperature and very high velocity oxygen stream. In the CHOR, this thermal nozzle is combined with a mixing chamber where the hot oxygen rapidly entrains a hydrocarbon feed gas, and a catalyst chamber where a high activity catalyst converts the reaction mixture into synthesis gas. Because of the rapid mixing and high velocities CHOR can safely handle flammable oxygen-hydrocarbon feed mixtures and operate at reactive conditions that have not been explored previously because of safety limitations. The development and performance of this system will be discussed in this paper. First a detailed description of the pilot plant used for this study is given followed by a description of the basic CHOR design. The hot oxygen mixing technique is described next followed by system description and performance results at low pressure. Then the system design and results for a high-pressure reactor are presented and finally results of the process economic analysis for the integration of CHOR with a gas-to-liquids process are discussed.

Figure 1. Catalytic hot oxygen reactor schematic.

tors. The system also includes atmospheric monitors, water flow switches, and the thermal oxidizer controls. A dual programmable logic controller (PLC) is used to monitor the safety interlocks, and a separate computer system is used to control and record the process variables. All data from process variables and analyzers are automatically saved in a spreadsheet to facilitate analysis. The reactor system is in an enclosure that is separated from the operator’s control room by a steel shield wall. This highly automated system allows for safe and precise experiments at a pilot scale. The carbon and H2 mass balances were calculated from the inlet gas flows and dry gas composition analysis of the product stream. O2 was used as the reference compound. H2 selectivity is calculated by selectivityH2 )

[H2] [H2] + [H2O]

(1)

where [C] denotes concentration in mol/cm3. Selectivity to CO is calculated by selectivityCO )

2. Pilot Plant Design The high-pressure experimental results reported here were produced in a pilot plant that was designed specifically to study catalytic partial oxidation reactions. The low-pressure reactor results were produced in an earlier facility that did not have all the automation features of the high-pressure pilot and also did not have online composition analysis but was similar in design. The low-pressure system will not be discussed here. A schematic of the pilot plant process flow diagram is presented in Figure 2. Oxygen and natural gas are compressed with air driven reciprocating compressors. All components wetted by pure oxygen were cleaned for oxygen service. The thermal nozzle fuel is supplied by a cylinder since its pressure is approximately twice that of the natural gas that is supplied to the mixing chamber. The natural gas is supplied by pipeline and is not desulfurized. All gas lines have low- and high-pressure switches, mass flow controllers, and pneumatic on/off valves. Pressure transducers and thermocouples near the reactor provide information on process pressure and temperature, respectively. A high temperature pressure control valve on the product stream is used to regulate reactor pressure. The reactor product is sent to a thermal oxidizer to convert synthesis gas to carbon dioxide and water before releasing them to a vent stack. A small portion of the product stream is sent to a five gas analyzer equipped with CO, H2, CO2, CH4, and O2 detectors and a gas chromatograph equipped with thermal conductivity and flame ionization detec-

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[CO] [CO] + [CO2] +

∑ [C ]

(2)

i

i

where Ci denotes the concentration of hydrocarbons produced in the reactor and excludes methane. Conversions of methane and O2 are calculated by molar flows: conversionO2 )

mol(O2)in - mol(H2O, CO, CO2)out mol(O2)in (3)

conversionCH4 )

mol(CH4)in - mol(CH4)out mol(CH4)in

(4)

Foam monolith catalysts, pellet catalysts, and combinations of the two were used in this study. Ceramic foam monoliths (Vesuvius Corp.) were impregnated with Rh catalyst by the insipient wetness method.3 The monoliths were first dipped in a Rh-acetone-acetonate solution, then dried and finally calcined at 600 °C. The procedure was repeated 3-4 times until about 2-5% of Rh was coated on the ceramic monolith. The pellet catalyst was supplied by UOP and was used as received. 3. CHOR System Description CHOR can be visualized in terms of three main elements: (1) thermal O2 nozzle, (2) mixing chamber in which the hydrocarbon feed is entrained into the hot oxygen stream to form a reaction mixture, which is delivered at high-pressure,

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Figure 2. Pilot plant process flow diagram.

high temperature, and high velocity to the catalyst, and (3) catalytic reaction chamber. Figure 1 shows a schematic of the CHOR. The thermal nozzle section of the CHOR is shown on the left side of Figure 1. In this region, a small amount of fuel gas is burned in excess oxygen in a diffusion flame. A spark electrode is used to ignite the flame. The combustion of the fuel gas supplies the energy to heat the unreacted portion of the oxygen stream. The temperature of the hot oxygen stream (as high as 1650 °C) can be precisely controlled by the amount of fuel fed to the diffusion flame. The hot oxygen then is expanded through a nozzle to form a hot oxygen jet. For a given nozzle diameter the maximum velocity of the hot oxygen stream is proportional to the square route of the temperature of the hot gas. The high temperature is also important for fast mixing with the primary fuel as will be explained in section 4. The hot oxygen stream that is expanded through the nozzle is mixed into a mixing chamber with the primary fuel reactant, which is added at the base of the hot oxygen jet. The nozzle is sized so that sonic or near sonic velocities are achieved. The pressure ratio between the thermal nozzle and the mixing chamber needs to be at least 2 to achieve sonic velocities at the exit of the nozzle. The mixing between the hot oxygen and the feed gas is very rapid due to the density difference between the hot oxygen and the colder fuel gas, and the turbulence introduced by the oxygen jet expansion. In the reactor designs described in this paper the mixing chamber volume was equal or less the catalyst volume so the mixing residence times are comparable with the residence times on the catalyst. The reactive mixture is fed to the catalyst before any significant homogeneous reaction can occur. The reaction on the catalyst is very rapid and highly exothermic. The catalyst can be in the form of foam or honeycomb monoliths or as conventional pellets. The CHOR thermal nozzle provides for efficient catalyst ignition without the requirement of any external heating. Catalyst ignition is defined as the initiation of the

reaction on the catalyst surface. In catalytic partial oxidation catalyst ignition occurs when the catalyst temperature reaches 300-400 °C. The temperature of the reaction mixture (hot O2 and natural gas) depends on the temperature and amount of the hot O2, and the temperature and amount of the hydrocarbon that are fed to the mixing chamber. The heat that is carried by the hot O2 can effectively heat the reaction mixture at temperatures above the catalyst ignition temperature. 4. Hot Oxygen Mixing The key to the CHOR system’s performance is the hot oxygen nozzle. As long as the temperature of the hot oxygen effluent does not exceed 1650 °C, it has been demonstrated that the thermal nozzle does not require any external cooling. It should be noted that steel parts cannot withstand such temperatures particularly in an oxygen atmosphere. It was hypothesized that the surprising mechanical stability of the thermal nozzle is a result of the following. (1). low-heat transfer rate. Heat transfer from oxy-fuel combustion products to a surface can be very intense because of the presence of active species such as atomic hydrogen and oxygen that recombine on the surface with a high heat of reaction. However, if the temperature is kept below 1650 °C, the concentrations of such active species are small. Therefore, the heat transfer from a hot oxygen stream not exceeding 1650 °C to a surface is an order of magnitude lower than that for combustion products from a stochiometric oxy-fuel burner. (2) Cool oxygen boundary layer. There is a cooler boundary layer of oxygen at the perimeter of the hot oxygen stream adjacent to the nozzle surface. Thus, the nozzle surface does not contact the hot oxygen. (3) Small exposed heat transfer area. The total heat transfer of the nozzle is relatively small. A computational fluid dynamic model was prepared using the geometry of the thermal nozzle described in the previous

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Figure 3. Temperature map of the thermal nozzle combustion chamber (°F).

section to test these hypotheses. The model included detailed turbulent flow description and a simple chemical reaction model to describe the combustion reaction. The assumption was that reaction kinetics are fast and mixing is the limiting factor. This approach is believed to be accurate because of the excess of oxygen used and the high pressure in the thermal nozzle. The results of the CFD model are presented in Figure 3, which depicts the temperature map of the thermal nozzle’s combustion chamber. It can be observed that a cool oxygen layer (blue region) is positioned between the flame and the metal wall of the thermal nozzle. This layer has a temperature of about 150-300 °C which is in qualitative agreement with our experimental observations. More importantly it confirms our primary hypotheses for the thermal nozzle mechanical stability. Another important feature derived from the modeling is the shape of the flame which initially expands and then contracts before it passes through the converging nozzle. The contraction is caused by the acceleration of the flow as it passes through the converging nozzle and prevents the hot gas from contacting the metal wall which could have caused it to overheat. These attributes of the thermal nozzle are critical to form a stable hot oxygen jet at the thermal nozzle exit, which is critical for good mixing performance. Mixing in the CHOR is based on the entrainment rate of the expanding hot O2 jet. The hot gas that is produced from the thermal nozzle entrains the primary fuel which is added at the base of hot gas jet. The entrainment rate for nonisothermal jets was described by Ricou and Spalding9 by the following equation:

()

m x F ≈ m0 d0 F0

0.5

(5)

where m is the total jet mass flow rate at distance x, m0 is the mass flow rate of the jet at the nozzle exit, x is the distance from the nozzle exit, d0 is the nozzle diameter, F is the density of the gas that is entrained into the jet and, F0 is the density of the jet at the nozzle exit. From eq 5 it can be deduced that the entrainment rate is proportional to the square route of the ratio of the jet gas temperature to the entrained gas temperature (density ratio). If the jet gas is at the same temperature as the entrained gas, the ratio of densities (F/F0) in eq 5 reduces to a constant equal to the ratio of the molecular weights of the two gases. If however the jet gas is at a temperature that is higher than that of the entrained gas, the mixing is enhanced by a factor equal to the

Figure 4. Composition profiles at the exit of the mixing chamber.

square route of the ratio of the temperatures of the two gases. Faster mixing allows flammable gases to be mixed before the onset of homogeneous reactions. In eq 5 the entrainment rate is also inversely proportional to the nozzle diameter (d0). Smaller nozzle diameter increases the entrainment rate (see eq 5). The smallest diameter of the nozzle is determined by the sonic velocity limitation. Higher hot gas temperatures correspond to higher sonic velocities so a smaller nozzle diameter can be used to accommodate the same flow rate. The fast, entrainment-based mixing achievable with the thermal nozzle is a critical aspect of CHOR and its ability to mix oxygen and hydrocarbons under potentially flammable conditions without igniting the mixture before it reaches the catalytic monolith. Propane, which has a wider flammability envelope than methane, was selected as the reactant to test the mixing capabilities of the thermal nozzle. Figure 4 depicts an experimental concentration profile for the CHOR shown in Figure 1 at a cross section at the end of the mixing chamber. Zero represents the center of the cross section. In this example, 12 slpm of natural gas and 145 slpm of oxygen were routed to the hot oxygen nozzle and an additional 95 slpm of propane gas was routed to the mixing chamber. The concentration profile was measured at the exit of the mixing chamber with the use of a small sampling tube with an opening of 1.6 mm. The exit of the mixing chamber had a diameter of 2.5 cm. From Figure 4 it can be observed that the concentration profiles are nearly flat. Small deviations near the edge can be attributed to low resolution of the sampling tube.

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one. The recirculation zones are still present but occur closer to the jet origin and are significantly reduced in size. Combined with the uniform composition the uniform velocity creates ideal conditions for the feed mixture to enter the catalyst chamber and creates a uniform radial temperature profile on the catalyst itself. Radial temperature differences are detrimental for the catalyst structural stability and can degrade its performance. In addition by reducing the size of the jet recirculation zones the possibility of initiating and stabilizing homogeneous reactions in the mixing chamber is reduced. 5. Catalyst Configuration (Low Pressure)

Figure 5. Pressure profiles across the exit diameter of the mixing chamber: (a) empty pipe (b) with monolith at end of mixing chamber

The flat concentration profiles mean that the mixture has a homogeneous composition just before it reaches the catalyst. The total propane flow at the exit of the mixing chamber was equal to the flow at the entrance therefore no homogeneous reaction was observed in the mixing chamber. The CO2 and H2O observed are combustion products from the thermal nozzle. It should be noted that the temperature at the exit of the mixing chamber was about 400 °C and the composition of the oxygen/ propane mixture was inside the flammability envelope. Nevertheless, no ignition was observed in the mixing chamber. The hot oxygen jet is not the only reason for the flat concentration profile at the exit of the mixing chamber. The presence of the catalyst in close proximity to the mixing zone impacts the mixing performance as well. The mixing performance was further tested by performing hydrodynamic experiments with and without the presence of a monolith in the catalyst chamber. A Pitot tube was used to measure pressure profiles inside the expanding section of the hot O2 jet. The Pitot tube was located at the exit diameter of the mixing chamber (same location as with previous experiment) and measured the pressure differential of the gas in the mixing chamber with atmospheric pressure. A measurement with an empty catalyst chamber closely resembled behavior of an axisymmetric jet expanding inside a duct.10 A parabolic pressure profile was measured (as shown in Figure 5a) which indicated a parabolic velocity profile was present. This profile is characteristic of large recirculation zones formed at the base of the jet. The zero pressure differential indicates a stagnation point. The pressure drop profile is not symmetrical so the jet is off center which may be due to geometry imperfections or an effect of the size of the probe (1.6 mm) and the small size of the mixing chamber which had an exit diameter of about 2.5 cm. A different behavior was observed with a monolith present in the catalyst chamber (Figure 5b). The monolith introduced flow resistance and changed the velocity profile to a uniform

The low-pressure system was very similar to the system depicted in Figure 1. The catalyst chamber could accommodate either pellets or monoliths or combination of both. The catalyst chamber was internally insulated with hard cast alumina insulation to reduce heat losses. About 5 cm of ceramic insulation was present around the catalyst. The catalyst chamber was made of a 15 cm diameter and 30 cm long schedule 40 stainless steel pipe. Results for three different configurations, (a) foam catalytic monolith, (b) catalytic pellets, or (c) combination of monolith and pellets, were in general agreement with Panuccio and Schmidt11 who showed that pellets are more productive than monoliths for air-based octane oxidation. Specifically we found that the Pt-γ-Al2O3 pellet catalyst was the most active compared to Rh on foam monolith catalyst and Rh on alpha alumina pellet catalyst. However, during oxygen-based experiments with Ptγ-Al2O3 pellet catalyst it was observed that some of the catalyst was destroyed due to high temperature exposure. The use of pure oxygen and the absence of diluent (e.g., nitrogen or steam) results in temperatures near the front of the catalyst well above 1000 °C when the oxygen/hydrocarbon ratio is near stochiometric for production of synthesis gas. It is also known that above 900 °C, γ-Al2O3 undergoes a phase transformation from gamma to alpha phase.12 This transformation destroys the highly porous gamma structure and results in a catalyst with markedly reduced activity. To protect the very active but not sufficiently stable pellet catalyst a configuration with three different layers of catalyst was used. The first layer was a Rh impregnated foam monolith 2.5 cm diameter by 2.5 cm long, the second layer was an alpha alumina pellet catalyst impregnated with Rh and the third layer was the Pt-γ-Al2O3 pellet catalyst. The use of the more stable but less active catalyst near the reactor front is acceptable since the high temperatures that develop in that region should provide sufficiently fast kinetics. Oxygen is depleted fast near the catalyst front and reforming reactions take over as gases move down the catalyst bed resulting in a temperature drop.13 Placing the more active but less stable γ-Al2O3 pellet catalyst at a point in the catalyst bed beyond the temperature maximum should give superior performance without compromising its integrity. Tables 1 and 2 present experimental results for propane and natural gas, respectively, for the catalyst configuration that combines monoliths and pellets. For synthesis gas production from propane and natural gas, the stochiometric O2/propane ratio is 1.5 and oxygen/natural gas is 0.65, respectively. The experimental focus was on proving the operability of the system and not on detail catalyst performance. Propane was supplied by cylinders and natural gas by pipeline. The natural gas was used as received with no pretreatment to remove sulfur compounds. No online analysis was available for these experiments so sample cylinders (500 cm3 volume) were filled with the synthesis gas product and were subsequently analyzed by a

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Table 1. Results with Propane Reactant. The O2 to Propane Ratio Takes into Account the O2 that Was Used in the Thermal Nozzle conversion %

selectivity %

O2 flow (slpm)

C3H8 flow (slpm)

O2/C3H8 ratio

pressure (bar)

feed temp (K)

middle temp (K)

exit temp (K)

C3H8

O2

CO

H2

38.5 88 41.5 69.8 79.2 70.8 70.8

63.7 42.9 42.9 55 54.3 43.1 42.9

0.44 0.56 1.16 1.37 1.60 1.79 1.94

1.6 1.4 1.5 1.88 2.02 1.82 1.92

524 731.6 681 465 455 653 666

1110 1047 1119 1388 1420 1502 1341

770.8 858 919.7 899 999 794.7 954.7

42.86 71.00 96.70 97.20 99.70 100.0 100.0

100.0 100.0 100.0 100.0 100.0 100.0 100.0

67.50 78.70 72.0 81.70 88.70 89.50 80.10

48.30 66.90 92.70 93.00 95.00 90.60 79.10

Table 2. Results with Natural Gas Reactant. The Oxygen to Fuel Ratio Takes into Account the Oxygen That Was Used in the Hot Oxygen Nozzle conversion %

selectivity %

O2 (slpm)

natural gas (slpm)

O2/natural gas ratio

pressure (bar)

feed temp (K)

middle temp (K)

exit temp (K)

CH4

O2

CO

H2

44.8 47 50.3 54.6 53.8 53.3 58.5

97.7 93.9 94.3 93.9 91.6 89.7 94.3

0.37 0.44 0.50 0.53 0.54 0.55 0.59

1.07 1.34 1.20 1.34 1.24 1.24 1.20

457 744 377 783 475 484 371

1129 870 1064 1087 1081 1100 1256

1027 505 785 814 755 773 790

75.40 71.60 87.80 90.60 90.90 91.88 98.00

98.70 98.90 98.80 98.88 98.90 99.00 98.90

90.86 86.13 91.00 87.80 90.20 90.60 92.00

96.10 90.40 94.80 94.40 92.70 92.00 93.00

gas chromatograph after the end of the experiment. The sample cylinders were purged with nitrogen and evacuated prior to use. This type of analysis could have introduced 2-5% error in our results. For the propane experiments natural gas 0.9-4.7 slpm was used as fuel to the hot O2 nozzle. The catalyst first layer was a Rh-impregnated foam monolith 2.5 cm diameter by 2.5 cm long, the second layer was an alpha alumina pellet impregnated with Rh (2 cm3) and the third layer was the γ-Al2O3 pellet catalyst (10 cm3). Space velocity was 240000-400000 h-1. From Table 1 it can be observed that CO and H2 selectivity initially increase as O2/fuel ratio increases and approaches the stochiometric ratio and then as the O2/fuel ratio increases above 1.5 hydrogen and CO selectivities decline due to the formation of more complete combustion products (CO2 and H2O) and also acetylene. Cracking products were also observed from cracking of propane but their concentrations decreased at higher oxygen/ propane ratios (with the exception of acetylene). Oxygen-propane mixtures with oxygen/propane ratio higher than 0.666 are inside the flammability envelope. The reactor operated without an indication that the oxygen-propane mixture ignited prematurely in the mixing chamber. For natural gas catalytic partial oxidation, natural gas 1.4-4.7 slpm was used as fuel to the hot oxygen nozzle. The first catalyst layer was a Rh-impregnated foam monolith 2.5 cm diameter by 2.5 cm long, the second layer was an alpha alumina pellet impregnated with Rh (20 cm3) and the third layer was a γ-Al2O3 pellet impregnated with Pt catalyst (300 cm3). Space velocity was 25-27000 h-1. Conversions and selectivities are comparable to those reported by Reyes et al.7 for a laboratory reactor operating with pure oxygen that was externally heated to minimize heat losses. In this work the ceramic insulation around the catalyst coupled with the higher flows used reduced the heat losses, and the reactor operated nearly adiabatically. One notable difference is that the hydrogen selectivity observed in this work is higher than the CO selectivity while the opposite is true for the laboratory system.7 This difference is likely due to the different catalyst configuration used in this work. No ignition was observed in the mixing chamber. Several continuous runs each lasting for over eight hours were conducted to test longterm performance of the system. The product composition remained virtually unchanged for the duration of these experi-

Figure 6. High-pressure reactor assembly drawing.

ments. No coke formation could be detected with visual inspection of the catalyst. 6. High-Pressure System 6.1. High-Pressure Reactor Design. The transition from low-pressure to high-pressure operation necessitated a change in the design of the reactor system for safety and operability reasons. The high-pressure reactor drawing is shown in Figure 6. The reactor was designed for operating up to 20 bar and the thermal nozzle for operating up to 40 bar. All reactor parts were constructed of stainless steel following ASME code specifications. The hot oxygen nozzle combustion chamber needed to be smaller since at higher pressures combustion reactions are faster and flame lengths decrease. No reliable expressions for predicting flame length are available so the system was designed with semiempirical expressions and tested at pressures up to 40 bar before being used in the reactor. The hot oxygen nozzle was water-cooled as a precaution to preserve its strength in case of overheating. The hot oxygen nozzle was water pressure tested

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Figure 7. Temperature and pressure profiles with original nozzle. The temperature increase in the front of the catalyst (TC4 and TC5) indicates homogeneous combustion.

at 750 bar. The catalyst cavity was 5 cm in diameter and 10 cm long. The catalyst section was internally insulated with about 10 cm of hard-cast ceramic insulation to protect the metal from overheating and minimize heat leaks. A stainless steel reactor shell supported the insulation and catalyst chamber and provided mechanical strength. The stainless steel reactor shell was pressure tested to 100 bar prior to use in experiments. The catalyst was placed at the center of the catalyst chamber and had a 2.5 cm diameter and was 5 cm long. The catalyst was immediately preceded and followed by uncoated alumina monoliths (2.5 cm diameter by 2.5 cm long each) that acted as radiation heat shields. The inert and active monoliths were wrapped with a high temperature alumina paper to minimize bypass of the reactant gases between the monolith and the reactor wall. High temperature alumina paper and graphite gaskets were placed between the mixing chamber and the catalyst chamber for better thermal insulation and sealing. The mixing chamber was water-cooled externally. 6.2. High-Pressure Reactor Mixing Chamber Design. For low-pressure experiments the thermal nozzle/mixing chamber design (Figure 1) was found to give excellent results. However, when it was attempted to go to higher pressures with the highpressure system the onset of homogeneous reactions occurred. High temperatures were observed well above the 150-200 °C adiabatic mixing temperature predicted in the absence of homogeneous reactions (Figure 7). In these experiments no catalyst was present in the catalyst chamber. From Figure 7 it can be observed that at pressure of about 50 psig thermocouple four (TC4) reported a high temperature spike from 200 °C to >900 °C indicating homogeneous reaction in the mixing chamber. To understand the problem computational fluid dynamic simulations (using CFX) were performed at low and high pressures. A comparison is depicted in Figure 8 where the composition map of O2 and methane inside of the mixing chamber is plotted for a low-pressure and high-pressure case, respectively. From Figure 8b the concentration profile shows that a larger recirculation zone forms at the base of the O2 jet at high pressure. The recirculation is not as large at low pressure as shown in Figure 8a. In the recirculation zone, hot gases slow down and can potentially ignite and stabilize a flame if the concentration of methane is inside the flammability envelope. At low pressure the recirculation zone is smaller and confined in the expansion cone of the mixing chamber. The geometry of the thermal nozzle exit was altered in order to reduce the open space before the mixing chamber. In doing so the velocity of the methane increases and the recirculation zones are pushed back into the expansion cone of the mixing

Figure 8. CFD simulation of mixing chamber (a) 1 bar and (b) 10 bar.

Figure 9. Schematic of the redesigned CHOR nozzle.

chamber. Such a design is depicted in Figure 9. An experimental nozzle was built as depicted in Figure 9 and tested with no catalyst present in the catalyst chamber. The mixing performance was stable and at 12 bar the temperature in the mixing chamber remained stable at about 175 °C. 6.3. High-Pressure Catalyst Bed Design and Results. The new nozzle configuration was subsequently tested with a catalyst present in the catalyst chamber. Reactive results with the new mixing chamber are presented in Figure 10. These experiments were performed with feeding the thermal nozzle with 6 slpm natural gas and 100 slpm oxygen and an additional 200 slpm of natural gas were fed in the mixing chamber. Typically after

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Figure 12. Schematic of the redesigned catalyst bed. Figure 10. Temperature and pressure profiles with redesigned nozzle. As pressure increases temperatures in the front of the catalyst (thermocouples TC4 and TC5) remain constant.

Figure 11. Temperature profiles along the catalyst bed for low and high pressure. The temperature maximum location moves closer to the front at higher pressure.

Figure 13. Temperature and pressure profiles with redesigned catalyst bed. As pressure increases no ignition occurs in front of the catalyst as indicated by thermocouple TC4.

the experiment starts and the catalyst is ignited the flow rates are fixed and the reactor pressure is increased by closing the back pressure regulator valve. Figure 10 shows the evolution of temperature at the exit of the mixing chamber and pressure at the catalyst chamber as a function of elapsed time from the start of the experiment. It is observed that as pressure increased to 150 psig (10 bar) the temperature recorded at thermocouple TC6 (Figure 6) increased from 350 to 900 °C. The temperature upstream of the catalyst was maintained constant up to that point (80 min). It is also observed that as thermocouple TC6 reaches 900 °C, the temperature at thermocouple TC4 and TC5 dramatically increased. This temperature rise indicates initiation of homogeneous reactions in the region before the catalyst. Since the problem was only observed with the catalyst present (see previous paragraph for results without a catalyst) and given the rise in temperature observed with thermocouple TC6, it was speculated that ignition is related to the catalyst presence. To elucidate the catalyst behavior a theoretical model developed by Vlachos et al.14 was used. The model qualitatively described the behavior at different pressures. The same operating conditions as those in the experiment were used in the model at two different pressures. Figure 11 presents the model predicted temperature profile inside the catalyst bed for two different pressures. The temperature profile exhibits a maximum near the front of the catalyst and this maximum moves closer to the front at higher pressures. This result explains the experimentally observed catalyst behavior at high pressures and agrees with experimentally observed temperature profiles at low

pressures.13 Because of the move of the temperature maximum at high pressure, the front of the catalyst reached high temperatures that initiated homogeneous reactions in front of the catalyst. The experimental data in Figure 10 also corroborate this hypothesis by showing that the temperatures in the front of the catalyst and before the catalyst increased with pressure. To overcome the reactant ignition before the catalyst at high pressure the temperature maximum must be prevented from moving toward the front of the bed. The following list summarizes possible solutions to this problem: (1) increase the velocity at the front of the catalyst; (2) remove some heat from the front of the catalyst as in a heat-integrated wall reactor;15 (3) operate with steam diluent to reduce the activity of the gas; (4) use a less-active catalyst for the first front of the catalyst. Only the results with the first approach were tested and are presented here. The catalyst bed design that was used is depicted in Figure 12 and was fitted in the catalyst cavity of the reactor described in Figure 6. The front of the bed had a narrower diameter than the back. A catalytic monolith 1.25 cm diameter by 2.5 cm long was used followed by an expanding diameter bed of catalytic pellets (section 5). The gas linear velocity through the front narrow section was higher than the wider back end. Since most of the O2 reacts in the front of the bed only the front needs to have higher velocity to prevent early occurrence of the temperature maximum. As the gas reacted it expanded due to the temperature increase and the bed design accommodated this expansion. The experimental results under reactive

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Figure 14. Gas to liquids process flow diagram. Table 3. Representative Operating Conditions at Different Pressures experiment 1 2 3

mixing Ch. temperature (°C)

catalyst front temperature (°C)

catalyst exit temperature (°C)

reactor pressure bar

flow (fuel natural gas, slpm)

flow (O2, slpm)

flow (natural gas, slpm)

213.6 244.4 230.2

514.6 724.5 741.7

989.0 978.1 1145.1

3.2 11.7 13.8

4.2 4.2 4.3

87.4 86.6 99.7

200.3 200.1 199.9

Table 4. Representative Synthesis Gas Composition at Table 3 Conditions experiment 1 2 3

H2

N2

0.5083 0.5040 0.5383

0.0187 0.0174 0.0172

synthesis gas composition mole fractions dry basis CO CO2 CH4 C2H6 0.2853 0.2764 0.2919

conditions with the new catalyst bed design are presented in Figure 13. The operating conditions are presented in Table 3. From Figure 13 it can be observed that the temperature in the mixing chamber remains around 250 °C while the reactor pressure increases to about 19 bar. No ignition occurred in the mixing area or in front of the catalyst. Table 4 presents the reactor performance results. Although significant hydrogen and carbon monoxide are produced some of the methane remains unreacted at the exit of the reactor. Reactor exit temperatures are also higher than those observed at low pressure. Lower space velocities should improve the quality of the product gas since the high temperatures recorded at the reactor exit indicate more heat is available to drive the conversion of methane to produce more CO and H2. 7. Process Economic Analysis The development of the high-pressure reactor was a promising success but the study of the integration of the new system with the gas-to-liquids process was also important to determine its commercial viability. The process and economic assessment was based on comparing a baseline process configuration for the production of liquid fuel from natural gas with one based on

0.0285 0.0321 0.0351

0.1572 0.1665 0.1124

0.0016 0.0017 0.0015

C2H4

C2H2

0.0005 0.0019 0.0033

0.0000 0.0001 0.0004

CHOR technology. This work was performed by Foster Wheeler. The “baseline” process configuration was for a standalone facility to produce 25 000 barrel per day (BPD) of distillates from natural gas. The baseline process configuration selected was representative of state-of-the-art commercially available synthesis gas production technology, and as such should be an effective yardstick against which the CHOR synthesis gas technology can be measured. Evaluation of the baseline configuration can also yield important clues about how CHOR can be used to best advantage in this application. The baseline configuration as shown in Figure 14 includes the following process steps: • autothermal reforming (ATR) of natural gas with oxygen and steam, at a low (0.6:1) steam to carbon ratio, to produce synthesis gas with a slight excess of hydrogen • Fischer-Tropsch (F-T) slurry-phase reactor to convert the synthesis gas into paraffins and olefins, primarily C20+ waxes • recycle of unconverted synthesis gas to the ATR and/or the F-T reactors • recovery of excess hydrogen with gas separation membranes • hydrotreatment of the F-T product waxes into transportation

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quality diesel, along with some lighter naphtha and kerosenerange distillates • electric power generation with a gas turbine, utilizing excess fuel gas • oxygen production via cryogenic air separation. Cost and economic estimates were developed of the two cases. One case was based on a conventional autothermal reformer and the other case was based on a CHOR reactor. Process flow schemes were developed, using literature information for F-T and hydro-isomerization sections with no licensor input. Factorial capital estimates were developed for both cases. Here are the assumptions of the economic analysis which was performed in 2003 so it may not include the latest state of the art: • generic marine tropical location (location factor 115%) • natural gas, 250 million standard cubic feet • maintenance, 3% inside battery limits (ISBL) capital • plant overhead, 70% direct fixed cost • insurance/property taxes, 1% total plant capital • environmental, 0.5% total plant capital • return on investment, 10% of ISBL and 5% outside battery limits (OSBL) • depreciation, 10% total capital investment As a result of the economic analysis it was concluded that CHOR had $2.30/bbl or 9% overall cost advantage relative to base case. The benefit results from a lower total capital cost (10.6% lower), reduced steam use (20% less steam), reduced O2 use (4% less O2), and increased production of liquids (2.% more liquids). CHOR process thermal efficiency was 62.3% versus 61.2% for the base case (LHV basis). 8. Discussion Process intensification by employing fast reaction kinetics has received considerable attention promising to reduce footprint and improve efficiency of chemical operations. Scale up and application of industrially relevant process conditions can result in surprising departures from bench scale performance. Studying the integration of the reactor with the rest of the plant and developing detailed process economics in parallel with the system development are necessary. A multidisciplinary team was assembled to investigate the possibility of applying short contact time reactor technology for the production of synthesis gas and its application to the production of liquids fuels. With proper mixer and catalyst design, a high-pressure pilot scale system was developed capable of producing synthesis gas from natural gas and pure oxygen. The system was tested at a pilot scale facility at the Praxair Technology Laboratory. A novel oxygen/fuel mixing device based on hot jet mixing was proven successful in mixing oxygen and fuel very close to the catalyst section. This mixing device is simple with a tube in tube configuration and is expected to scale up to larger sizes. It can achieve mixing in times that are comparable to the reaction times on the catalytic monoliths so a compact reactor can be constructed. One disadvantage of the system is the requirement of high-pressure oxygen (twice the pressure of the reactor chamber). At high-pressure operation the reactants linear velocity near the top of the catalyst section must be sufficiently high to prevent the temperature maximum from occurring near the top of the bed since it could lead to homogeneous reactions before the catalyst bed. Temperatures observed near the top of the bed with the use of pure oxygen as the oxidant were high and could be detrimental for the catalyst stability. A layered catalyst configuration with a less active but more stable catalyst at the top of the catalyst bed, followed by subsequent layers of less stable but more active catalyst was a

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good fit for this system. The high temperatures that develop near the top of the catalyst where oxygen is consumed aid reaction kinetics that should be sufficiently fast even with a less active catalyst. A less stable but more active catalyst can be used as temperatures drop and reforming reactions dominate downstream of the catalyst bed. Longer tests than those reported here are required to verify catalyst life particularly at high pressures. The economic analysis for the production of liquid fuels from natural gas favored a catalytic oxidation system over conventional technology but the advantage was small. The application of this system to other catalytic reactions will be explored in the future. Acknowledgment I would like to express my gratitude for the efforts of Foster Wheeler and U. of Delaware without which this work would not have been competed. I would also like to express my gratitude for National Energy & Technology Laboratory who provided help for this work through cooperative agreement DEFC26-00NT41027. V.P. This paper was prepared with the support of the U.S. Department of Energy, under Award No. DE-FC26-00NT41027; however, any opinions, findings, conclusions, or recommendations expressed herein are those of the authors and do not necessarily reflect the views of the DOE. Literature Cited (1) Gunarson H. Industrial Gases In Petrochemical Processes; Marcel Dekker Inc.: 1997. (2) Naqvi S. N. Opportunities for Gas-To-Liquids Technologies. Process Economics Program Report 135C; SRI Consulting: Menlo Park, CA, 2000. (3) Hickman, D. A.; Schmidt, L. D. Synthesis Gas Formation by Direct Oxidation of Methane over Pt Monoliths. J. Catal. 1992, 138, 267. (4) Hickman, D. A.; Schmidt, L. D. Steps in CH4 Oxidation on Pt and Rh Surfaces: High-Temperature Reactor Simulations. AIChE J. 1993, 39, 1164. (5) Bodke, A. S.; Bharadwaj, S. S.; Schmidt, L. D. The Effect of Ceramic Supports on Partial Oxidation of Hydrocarbons over Noble Metal Coated Monoliths. J. Catal. 1998, 179, 138. (6) Van den Sype J. S.; Barlow A. R. Process and Apparatus for Producing Heat Treating Atmospheres. U.S. Patent No. 5,441,581, 1995. (7) Reyes, S. C.; Sinfelt, J. H.; Feeley, J. S. Evolution of Process for Synthesis Gas Production: Recent Developments in an Old Technology. Ind. Eng. Chem. Res. 2003, 42, 1599. (8) Anderson J. E. Thermal Nozzle Combustion Method. U.S. Patent No. 5,266,024, 1993. (9) Ricou, F. P.; Spalding, D. B. Measurements of Entrainment by Axisymmetrical Turbulent Jets. J. Fluid Mech. 1961, 11, 21. (10) Beer J. M.; Chigier N. A. Combustion Aerodynamics; Robert E. Krieger Publishing: Malabar, FL, 1983. (11) Panuccio, G. J; Schmidt, L. D. Species and Temperature Profiles in a Differential Sphere Bed Reactor for the Catalytic Partial Oxidation of n-Octane. Appl. Catal., A. 2007, 332, 171. (12) Chai, M.; Machida, M.; Eguchi, K.; Arai, H. Preparation and Characterization of Sol-Gel Derived Microporous Membranes With High Thermal Stability. J. Membr. Sci. 1994, 96, 205. (13) Horn, R.; Degenstein, N. J.; Williams, K. A.; Schmidt, L. D. Spatial and Temporal Profiles in Millisecond Partial Oxidation Processes. Catal. Lett. 2006, 110, 169. (14) Park, Y. K.; Aghalayam, P.; Vlachos, D. G. A Generalized Approach for Predicting Coverage-Dependent Reaction Parameters of Complex Surface Reactions: Application to H2 Oxidation Over Platinum. J. Phys. Chem. A 1999, 103, 8101. (15) Piga, A.; Ioannides, T.; Verykios, X. E. Synthesis Gas Formation by Catalytic Partial Oxidation of Methane in a Heat-Integrated Wall Reactor . Studies in Surface Science and Catalysis; Natural Gas Conversion V, vol. 119; Elsevier: New York, 1998; p 411.

ReceiVed for reView May 27, 2009 ReVised manuscript receiVed November 2, 2009 Accepted November 4, 2009 IE900872R