Synthesis Gas Reformers

July 9, 2017 | Author: Albert Powa | Category: Heat, Carbon, Thermodynamic Equilibrium, Chemical Reactions, Catalysis
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SYNTHESIS GAS REFORMERS Megan Strait, Glenda Allum, Nisha Gidwani TABLE OF CONTENTS ABSTRACT TABLE OF CONTENTS 1.0 INTRODUCTION 2.0 OPTIMUM PROCESS 2.1 STEAM/CARBON SENSITIVITY 2.2 OXYGEN SENSITIVITY 2.3 PRESSURE SENSITIVITY 2.4 TEMPERATURE SENSITIVITY 3.0 REACTOR SIZING 4.0 PROCESS ECONOMICS 5.0 CONCLUSIONS AND RECOMMENDATIONS REFERENCES 1.0 INTRODUCTION Primary and secondary reformers play an important role in the production of ammonia. The goal of reforming is to prepare as pure as possible a gas mixture of nitrogen and hydrogen in a 3:1 stoichiometric ratio from the raw materials of water, air, and natural gas. The reactions by which this ratio are achieved are given as follows:

Reaction 1, the steam reforming reaction, and reaction 2, the water gas shift reaction, are endothermic and occur in the primary reformer. Reaction 3, the combustion reaction, is exothermic and occurs along with reactions 1 and 2 in the secondary reformer [1]. Optimization of the reforming process involves the manipulation of parameters to achieve high process yield while maintaining low operating and installed costs. The parameters which are monitered in this design include temperature, pressure, steam to carbon ratio, and percent oxyegen in the air feed.

2.0 OPTIMUM PROCESS The focus of this project is on the design of the primary and secondary reformers. There are, however, several other reactors necessary in any ammonia production process. These reactors can be seen in the Process Flow Diagram (PFD) in Figure 1. A brief overview of these other processes as well as the assumptions NGM Reformers made in considering the effects of these processes on the primary and secondary reformers will be discussed here. As seen in Figure 1, ammonia production begins with inlet streams of natural gas and steam. For the optimum process, NGM chose a 20225.902 kg/hr natural gas feed stream composed of 99% methane entering at a pressure of 35.29 bar. A 3:1 molar steam to carbon ratio was chosen for the feed streams. The steam stream enters with a flowrate of 67557.297 kg/hr [2]. The methane feed is first passed through a desulfurizer and then mixed with the steam feed. This mixed stream, the primary reformer inlet stream of Table 1, next passes through a heat exchanger where it is heated to a temperature of 600 C. From here the primary reformer inlet stream enters the primary reformer. The primary reformer is modeled in ASPEN as an RGibbs reactor with RK Soave equations of state. An operating temperature of 800 C and an operating pressure of 35.29 bar are found to yield the optimum process based on an analysis of sensitivity runs and industry standards. The primary reformer is heat flux limited. Thus, based on a calculated reformer heat duty of 50 MMkcal/hr, the reformer is sized to contain 230 catalyst tubes with 4" inner diameter and a length of 35'. From the primary reformer, the primary outlet stream is mixed with a 600 C, 35.29 bar air stream. This mixed stream constitutes the secondary inlet stream in Table 1. The secondary reformer is also modeled in ASPEN as an RGibbs reactor with RK Soave equations of state. An operating temperature of 996.2 C and operating pressure of 35.29 bar are found to best optimize the process. Based on these process conditions the secondary outlet stream flows of hydrogen, carbon monoxide and nitrogen were found to be 3375.48 kmol/hr, 814.378 kmol/hr, and 1350.593 kmol/hr respectively. These numbers are critical to the production of 1000 metric tons per day of ammonia in the correct 3:1 stoichiometric ratio. Based on the assumption that all downstream processes are ideal, all the carbon monoxide found in the secondary reformer outlet stream will be converted to hydrogen. Therefore, the addition of the hydrogen and carbon monoxide should be in at least a 3:1 ratio with nitrogen upon exiting the secondary reformer. This result is accomplished and is evident in Table 1. From the secondary reformer, the secondary effluent flows through a heat exchanger where it heats the primary reformer inlet stream. From the heat

exchanger it passes through a high temperature shift converter, low temperature shift converter, and methanator. For this projec these processes were assumed to operate ideally at equilibrium. FIGURE 1:


2.1 STEAM/ CARBON SENSITIVITY Primary reformer inlet steam-to-carbon (s/c) ratio is an important factor in reformer design. The literature advises the maintenance of a relatively high s/c ratio to prevent mechanical as well as economic problems during the life of the plant. Higher s/c ratios are more effective for a number of reasons. First, because a high s/c ratio favors the products in the reforming reaction equilibrium, it lowers the amount of unreacted methane, or methane slip, out of the secondary reformer and increases the production of hydrogen. Second, a high s/c ratio inhibits the occurence of carbon-forming side reactions in the primary reformer that result in carbon deposits on the catalyst. Carbon deposition increases the resistance to gas flow in the primary reformer tubes and may impair catalyst activity. This impairment lowers the rate of the reforming reaction and can cause local overheating or "hot bands" in reformer tubes that result in premature tube wall failure. Finally, a high s/c ratio provides the necessary steam for the shift conversion of carbon monoxide and reduces the risk of carburization damage to the tube material [3]. Table 2 shows process sensitivity to s/c ratio. A 3.0/1.0 s/c ratio was found to be the most optimum ratio for the purposes of this process. Sensitivity runs in Aspen showed that 4.0/1.0 s/c ratio requires a larger heat duty than a 3.0/1.0 ratio. This increases cost as more heat has to be applied to the process. However, lowering the s/c ratio to 2.5/1.0 was found to increase methane slip significantly, decreasing the amount of hydrogen produced. For these reasons a 3.0/1.0 s/c ratio was chosen.


2.2 OXYGEN SENSITIVITY Oxygen-enriched air is sometimes utilized in the production of syn-gas as it shifts more of the reforming from the primary reformer to the secondary reformer. An increase in the proportion of reforming occurring in the secondary reformer results in a higher outlet temperature from the secondary reformer. This heat can be recycled and used to heat the primary reformer inlet stream to reduce energy costs. On the other hand, enriched air introduces another cost to the process by requiring that excess nitrogen be stripped from the process downstream or that excess oxygen be purchased from a third party supplier [4]. NGM Reformers Inc. decided that this extra cost exceeded the savings gained from reducing energy costs. Hence, pure air (21% oxygen, 78% nitrogen, 1% inerts) is used. Using enriched air also decreases methane slip considerably, however, it decreases the production of hydrogen to below product specifications. Table 3 shows process sensitivity to the use of an enriched air stream. TABLE 3

2.3 PRESSURE SENSITIVITY The chemistry, economics and demands from major clients must be taken into consideration when analyzing the effects of reactor operating pressures on the process. It is important to note that the process is limited by a maximum pressure of 40 bar due to the metallurgy of the material used to construct the primary reformer tubes. Table 4 demonstrates the results of raising and lowering reformer pressure. High reformer pressures near 40 bar favor the reactants of the reforming

reaction equilibrium, therefore, the production of hydrogen decreases while methane slip increases. To compensate for high methane slip the heat duty must be increased, thus increasing compression and energy costs. Higher pressures also cause the secondary reformer effluent temperature to decrease. This has an unfavorable effect on the process as the heat from this stream is used to heat the primary reformer inlet stream via heat exchanger. A lower pressure of 25 bar exerts a favorable effect on the equilibrium of the reforming process. According to the chemistry involved it seems that lower pressures afford the most advantage; they increase secondary reformer outlet temperature, decrease methane slip to about 0.01%, and increase hydrogen production by approximately 100 kmol/hr. NGM Reformers, Inc. values customers in downstream ammonia synthesis and realizes that their compression costs will be significantly increased by low presures in the front end process. Hence, a process pressure of 35.29 bar was adopted as a fair compromise, allowing for the maintenance of a relatively low methane slip of 0.35% and a sufficient throughput of product while maintaining a high enough temperature in the secondary reformer outlet stream TABLE 4

2.4 TEMPERATURE SENSITIVITY The reforming process favors high temperatures as it shifts the reforming reaction equilibrium towards the production of hydrogen and reduces methane slip. However, it is not advisable to operate the primary reformer above 800 C because the metallurgy of the catalyst tubes causes them to creep and bulge under the weight of the catalyst at approximately 850 C. Additionally, the nickel catalyst melts at 1100 C. Operating at elevated temperatures also increases the heat duty, causing energy costs as well as equipment costs to escalate somewhat. In contrast, operating at 700 C decreases hydrogen production and increases methane slip out of the secondary reformer resulting in the waste of fuel. Table 5 shows these effects. The only advantage of lowering temperature is a decrease in heat duty, which will reduce costs somewhat. Because of the large degree of process senstivity to primary reformer temperature, it is desirable to operate at a temperature as close to the metallurgical limit of 850

C as possible in order to maximize H2 production. Therefore a temperature of 800 C in the primary reformer was chosen. This temperature is transferred to the adiabatic secondary reformer. The secondary reformer is not constrained by mechanical heat transfer surfaces; therefore, it can operate at higher temperatures and is operated at 966 C outlet temperature. TABLE 5

3.0 REACTOR SIZING Primary and secondary reformer reactor sizes were calculated from industry data in order to minimize primary reformer size and thus minimize installed cost. The primary reformer is heat flux limited; that is, reactor size is determined based on the surface area over which the necessary heat for reforming is transferred. A tube size of 4 inch I.D., 35 feet in length was chosen. This tube size is consistent with industry averages [5]. The maximum conventional heat flux through primary reformer tube walls is approximately 21,000 Btu/ft2*hr (5,921.176 kcal/ft2*hr) [6]. Using this value and the heat duty through the reformer calculated by Aspen, the primary reformer size was calculated as follows: f = Maximum heat flux thorough tube walls = 5,921.176 kcal/ft2*hr d = Heat duty through primary reformer (from Aspen) = 50.0771 x 106 kcal/hr a = Total needed surface area of reformer tubes = d/f = 8457.28 ft2 t = a/36.7 ft2 per tube = 230 tubes needed Catalyst volume was calculated from tube number and tube volume. The primary reformer contains a total of 690 ft3 of catalyst. The secondary reformer size was chosen based on industry input and a length to diameter ratio of approximately one [5]. The reactor is 12 feet in diameter and 20 feet long. Ten feet of reactor length are left void of catalyst so that combustion may occur away from the catalyst.

4.0 PROCESS ECONOMICS Approximate price ranges were obtained from M.W. Kellogg for the primary and secondary reformers and catalyst. These prices reflect the differences in construction materials used for each reactor. Due to the high pressures and temperatures in the primary reformer tubes, a 25% chromium-20% nickel alloy is the preferred tube material. The secondary reformer, with its simpler design, can be priced as a large, refractory-lined vessel containing a fixed-bed nickel catalyst [7]. Primary reformers cost on the order of $5 million, secondary reformers on the order of $1 million, and primary reformer catalyst approximately $200/ft3 [5]. Therefore catalyst cost for the primary reformer is $138,000, less than three percent of the total primary reformer installed cost. Since the primary refomer is such a major component of the process cost, the process was optimized so as to minimize the size of the primary reformer. Less attention was given to the amount of catalyst supplied to the primary reformer since it becomes almost negligible when compared to the cost of the reactor itself. 5.0 CONCLUSIONS AND RECOMMENDATIONS In this process a 3:1 stoichiometric ratio of (H2 + CO) to N2 in the secondary reformer effluent was achieved using a 90,000 kg/hr steam plus natural gas basis feed with 3:1 steam to carbon ratio, 35.29 bar reformer operating pressures, and 800 and 996 degrees C operating temperatures in the primary and secondary reformers, respectively. Several other factors may be considered, however, in designing a fully optimum process for the required syngas output. These include catalyst type, kinetic data, and reforming heat exchange. The process modeled and optimized here is an equilibrium model. That is, each reactor is assumed to be operating at equilibrium. In reality however, equilibrium can not be reached, and an approach to equilibrium model should be adopted. This involves the selection and incorporation of the approprate kinetic data for the process. The approach to equilibrium in the reformers is affected by choice of catalyst. Numerous catalysts of differing properties are available to suit specific purposes. Once a catalyst is chosen, parameters such as catalyst activity, surface area, particle size, crush strength, and nickel content among others should be considered to more accurately model the process. An alternative to the placement of the single heat exchanger used in this process is to employ reforming heat exchange. This involves recycling the secondary reformer effluent stream into the shell side of the primary reformer vessel such that the heat exchange between this stream and the primary reformer inlet occurs physically inside the primary reformer vessel rather than through a seperate heat exchange unit. This would require additional mechanical considerations and an alteration of the construction of the primary and secondary reforming vessels but is extremely thermodynamically favorable. Reforming heat exchange may

significantly reduce the size of the primary reforming furnace and consequently decrease the cost of the unit. Thus, reforming heat exchange should be considered as a possible alternative heat recovery mechanism [5,7]. REFERENCES [1]. Gerhartz, W. et. al. Ullmann's Encyclopedia of Industrial Chemistry. 5th Edition. VCH, Federal Republic of Germany: 1985. [2]. Mii, T. Process Systems Planning & Engineering Division, Toyo Engineering Corporation, Japan. Electronic mail correspondance. [3]. Appl, M. Modern Production Technologies. British Sulphur Publishing, London: 1997. [4]. Ned, Zed, and Associates and The Autobots. "Design and Economic Analysis of Ammonia Production Plant," CENG 404 Project Report, Spring 1996. [5]. Strait, R. B. Process Engineer, M.W. Kellogg Co., Houston, TX. Personal interview. [6]. Strelzoff, S. Technology and Manufacture of Ammonia. John Wiley & Sons, New York: 1981 [7]. Kirk-Othmer. Encyclopedia of Chemical Engineering Technology. 4th Edition. John Wiley & Sons, New York: 1992.

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