Campbell Tip of the Month

Author: Dr. Mahmood Moshfeghian

  • Effect of Impurities on Propane Refrigeration System – Constant Approach Temperature

    In this Tip of the Month (TOTM), we will continue our discussion of the effect of working fluid impurities on the performance of refrigeration systems employing propane as the working fluid. Specifically, we will study the effect on the compressor power, the refrigerant circulation rate, and the condenser duty.

    In the previous TOTM, the chiller inlet temperature was kept at -35°C but the chiller outlet temperature varied due to the presence of impurities; therefore, the approach temperature changed. In this TOTM, we will revisit the same case for constant approach temperature. In other words, the chiller outlet temperature was kept at -35°C but the chiller inlet temperature/and pressure varied due to the presence of impurities.

    The details of a simple single-stage refrigeration system and a two-stage refrigeration system employing one flash tank economizer are given in Chapter 16 of Gas Conditioning and Processing, Volume 2 [1]. Similar to the previous TOTM, the process flow diagrams for the simple and with flash economizer refrigeration systems are shown in Figure 1. Note that pressure drop in different segments of the loops have been considered.

    Figure 1

    Let’s consider the case of the previous TOTM in which the objective was to remove 2778 kW from the process gas at -35°C and rejecting it to the environment by the condenser at 35°C. The pressure drop assumptions were: in the line from the compressor discharge to the condenser and in the condenser 50 kPa, in the chiller 5 kPa and in the compressor suction line 30 kPa, between the two stages of flash economizer 20 kPa and between the flash tank and second stage of compressor 20 kPa. Pure propane was used as the base working fluid. An isentropic efficiency of 75 % was used in all cases. For the flash tank economizer, the optimized interstage pressure which minimized the total compressor power was used. In this study, all of the simulations were performed by HYSYS software [2].

    The production of propane is through fractionation of NGL; however, achieving a purity of 100% is not economical. Therefore, propane as the working fluid has normally a small fraction of ethane and butanes. We considered these components as impurities and studied their effect on the performance of the propane refrigeration system. The composition and molecular weight of the eleven cases studied are shown in Table 1. The last column represents the ratio of mixture molecular weight to the molecular weight of propane. As in the previous TOTM, we considered the single stage (simple) and the two stage (economizer) refrigeration systems. The summary of simulation results is shown in Tables 2 A&B.

    Tables 1 and 2A 

    Table 2B

    For graphical representation of the simulation results, the data in Tables 2 A&B are plotted in dimensionless form in Figures 2 through 4. Pure propane refrigeration system has been chosen as the base case and the performance of other cases are compared to the base case. Figure 2 represents the effect of ethane and butane impurities on the required circulation rate. Note that in this figure and in the subsequent figures, the y-axis is the ratio of case 2 through 11 variables (circulation rate, condenser duty or compressor power) divided by the corresponding case 1 variable, respectively.  Similarly, the x-axis is the ratio of cases 2 through 11 molecular weights to the molecular weight of case 1. Recall that case 1 is pure propane which was used as the base case.  As shown in this figure, the ethane impurity increases the circulation rate. The increasing butanes impurities cause a decrease in the circulation rate. The effect of ethane is approximately twice that of butane for a given impurity level.

    Figure 2

    Figures 3 and 4 represent the effect of ethane and butanes impurities on the condenser heat duty and the required compressor power requirement, respectively. These two figures indicate that both ethane and butane impurities increase the condenser duty and the compressor power requirement. It is interesting to note that the effect of butane impurities is two times higher than ethane impurities for the same level (mole fraction) of impurities.

    On reviewing Figures 2 through 4, the following observation can also be made:

    1. The impurities affect the performance of the simple refrigeration systems.
    2. The trend of impurity effect is similar for the simple refrigeration system and the refrigeration system employing a flash tank economizer.
    3. In order to keep the chiller outlet temperature at the specified value, the incoming refrigerant temperature had to be decreased which resulted in lower chiller pressure (See Tables 2 A&B). This caused an increase in the compression ratio and consequently in higher compressor power consumption. Also, as the ethane impurity increases, the compressor discharge (as well as condenser) pressure increases. In case of simple refrigeration, the compressor discharge pressure increases from 1270 to 1417 kPa when ethane mole fraction is changed from 0 to 5 percent.
    4. For the case of pure propane, the compressor power and condenser heat duty are minimum.
    5. For the economizer, the feed to the first stage of the compressor is heavier than the feed to the second stage due to the mechanisms of flash separation. The heavier components go with liquid stream and the lighter components go with vapor stream.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By Dr. M. Moshfeghian

    Figures 3 and 4

    Reference:

    1. Campbell, J.M., “Gas conditioning and Processing, Vol 2: The Equipment Modules”, 8th Edition, Edited by R.A. Hubbard, John M. Campbell & Company, Norman, USA, 2000.
    2. ASPENone, Engineering Suite, HYSYS Version 2006, Aspen Technology, Inc., Cambridge, Massachusetts U.S.A., 2006.

  • Effect of Impurities on Propane Refrigeration System

    In this Tip of the Month (TOTM), we will demonstrate the effect of working fluid impurities on the performance of a simple propane refrigeration system and another one employing a flash economizer. Specifically, we will study the effect on the compressor power, the refrigerant circulation rate, and the condenser duty.

    The objective of a refrigeration system is to “pump” low temperature heat from the process fluid to high temperature (ambient) where it is rejected to the environment. Energy is required to pump heat. The amount of energy depends on the quantity of heat to be “pumped’ (chiller duty) and how far the heat has to be pumped (temperature difference between the chiller and the condenser). Compression refrigeration is by far the most common mechanical refrigeration process. It has a wide range of applications in the gas processing industry. It provides chilling for:

    1. NGL extraction, LNG production, and LPG product storage
    2. Hydrocarbon and water dew point control
    3. Reflux in deethanizers/demethanizers

    The details of a simple single-stage refrigeration system and a two-stage refrigeration system employing one flash tank economizer are given in Chapter 16 of Gas Conditioning and Processing, Volume 2 [1]. Similar to the previous TOTM, the process flow diagrams for the simple and with flash economizer refrigeration systems are shown in Figure 1. Note that pressure drop in different segments of the loops have been considered.

    Process Flow Diagram

    Figure 1. Process flow diagram for the simple and flash economizer refrigeration systems

    Let’s consider the case of the previous TOTM in which the objective was to remove 2778 kW from the process gas at -35°C and rejecting it to the environment by the condenser at 35°C. The pressure drop assumptions were: in the line from the compressor discharge to the condenser and in the condenser 50 kPa, in the chiller 5 kPa and in the compressor suction line 30 kPa, between the two stages of flash economizer 20 kPa and between the flash tank and second stage of compressor 20 kPa. Pure propane was used as the base working fluid. An isentropic efficiency of 75 % was used in all cases. For the flash tank economizer, the optimized interstage pressure which minimized the total compressor power was used. In this study, all of the simulations were performed by HYSYS software [2].

    The production of propane is through fractionation of NGL; however, achieving a purity of 100% is not economical. Therefore, propane as the working fluid has normally a small fraction of ethane and butanes. We considered these components as impurities and studied their effect on the performance of the propane refrigeration system. The composition and molecular weight of the eleven cases studied are shown in Table 1. The last column represents the ratio of mixture molecular weight to the molecular weight of propane. As in the previous TOTM, we considered the single stage (simple) and the two stage (economizer) refrigeration systems. The summary of simulation results is shown in Table 2.

    Tables 1 and 2

    For graphical representation of the simulation results, the data in Table 2 are plotted in dimensionless form in Figures 2 through 4. Pure propane refrigeration system has been chosen as the base case and the performance of other cases are compared to the base case. Figure 2 represents the effect of ethane and butane impurities on the required circulation rate. Note that in this figure and in the subsequent figures, the y-axis is the ratio of case 2 through 11 variables (circulation rate, condenser duty or compressor power) divided by the corresponding case 1 variable, respectively.  Similarly, the x-axis is the ratio of cases 2 through 11 molecular weights to the molecular weight of case 1. It should be reminded that case 1 is pure propane which was used as the base case.  As shown in this figure, contrary to butanes, the ethane impurity has no practical effect the circulation rate. The increasing butanes impurities cause a decrease in the circulation rate.

    Figure 2

    Figures 3 and 4 represent the effect of ethane and butanes impurities on the condenser heat duty and the required compressor power requirement, respectively. These two figures indicate that butane impurities don’t have practical effect on the condenser duty and the compressor power requirement. However, ethane impurities increase the condenser duty and the compressor power requirement. In this study, the chiller inlet temperature was kept at -35°C but the chiller outlet temperature varied due to the presence of impurities; therefore, the approach temperature changed. In the next TOTM, we will revisit this case for constant approach temperature.

    On reviewing Figures 2 through 4, the following observation can also be made:

    1. The impurities affect the performance of the simple refrigeration systems.
    2. The trend of impurity effect is similar for the simple refrigeration system and the refrigeration system employing a flash tank economizer.
    3. The effect of impurities on the performance of a flash tank economizer is less pronounced than on a simple refrigeration system.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By Dr. M. Moshfeghian

    Figures 3 and 4

    Reference:

    1. Campbell, J.M., “Gas conditioning and Processing, Vol 2: The Equipment Modules”, 8th Edition, Edited by R.A. Hubbard, John M. Campbell & Company, Norman, USA, 2000.
    2. ASPENone, Engineering Suite, HYSYS Version 2006, Aspen Technology, Inc., Cambridge, Massachusetts U.S.A., 2006.

  • Refrigeration with Flash Economizer vs Simple Refrigeration System

    In this Tip of the Month, we will compare the performance of a simple refrigeration system with another one employing a flash economizer. Specifically, we will evaluate compressor power saving, the effects of compressor suction–line pressure drop and the interstage pressure drop on compressor power requirement and condenser duty.

    The objective of a refrigeration system is to “pump” low temperature heat from the process fluid to high temperature (ambient) where it is rejected to the environment. Energy is required to pump heat. The amount of energy depends on the quantity of heat to be “pumped’ (chiller duty) and how far the heat has to be pumped (temperature difference between the chiller and the condenser). Compression refrigeration is by far the most common mechanical refrigeration process. It has a wide range of applications in the gas processing industry. It provides chilling for:

    1. NGL extraction, LNG production, and LPG product storage
    2. Hydrocarbon and water dew point control
    3. Reflux in deethanizers/demethanizers

    The details of a simple single-stage refrigeration system and a two-stage refrigeration system employing one flash tank economizer are given in Chapter 16 of Gas Conditioning and Processing, Volume 2 [1]. The process flow diagrams for the simple and with flash economizer refrigeration systems are shown in Figure 1. Note that provisions have been made to consider pressure drop in different segments of the loops.

    Process Flow Diagram

    Figure 1. Process flow diagram for the simple and flash economizer refrigeration systems

    Let’s consider removing 1.0×107 kJ/h (2778 kW) from the process gas at -35°C and rejecting it to the environment by the condenser at 35°C. Assuming 5 kPa pressure drop in the chiller, the compressor suction pressure is 132.4 kPa. The condenser pressure drop plus the pressure drop in the line from the compressor discharge to the condenser was assumed to be 50 kPa; therefore, compressor discharge pressure is 1270 kPa. Pure propane is used as the working fluid. The effect of impurities in the working will be discussed in the future Tip of the Month. In this study, all of the simulations were performed by HYSYS software [2].

    For the case of with flash economizer and assuming no pressure drop between the two stages and in the suction line, Figure 2 presents the variation of the compressor interstage and total power as a function of the interstage pressure. As can be seen in this figure, the optimum interstage pressure is about 490 kPa. This pressure corresponds to the minimum total power. However, the ideal optimum interstage pressure based on equal compression ratio can be found by Equation kPa. Figure 3 also presents the locations of these two optimum pressures.

    Figures 2 and 3

    In order to study the effect of pressure drop in the compressor suction line on the total power requirement and condenser duty, the interstage pressure drop was set equal to zero. The suction line pressure drop was varied from 0 to 30 kPa with an increment of 10 kPa. Two sets of simulations were performed:

    1. The interstage pressure was determined based on the equality of compression ratio
    2. The interstage pressure was determined by minimizing the total compressor power

    In each case, the total compressor power for the flash economizer system was compared with the power requirement for the simple refrigeration system and the percent of power savings were calculated.

    Next, the effect of the interstage pressure drop on the total compressor power requirement and condenser duty was investigated. This was done by setting the compressor suction line pressure drop to 30 kPa and varying the interstage pressure drop from 0 to 40 kPa by an increment of 10 kPa. The load of interstage pressure drop was equally distributed between the two stages of compression. Again, the simulation results of flash economizer were compared with those of the simple refrigeration system. The summary of the results is shown in Table 1. Figures 4 and 5 are the graphical representation of the results presented in Table 1.

    Table 1. Comparison of the results based on the equality of compression ratio and minimizing the total power requirement

    Table 1

    Table 1 indicates that the power saving ranges from 12.7 to 16.1 % when a flash economizer is used in place of the simple refrigeration system, using interstage pressure based on the equality of compression ratio. However, if the interstage pressure is determined based on minimizing the total compressor power requirement, the power saving will be from 14.4 to 16.5 %. The interstage pressure drop is unique to flash economizer and its effect is the reduction of the power saving when compared to the simple refrigeration system and increases the condenser duty.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By Dr. M. Moshfeghian

    Figures 4 and 5

    Reference:

    1. Campbell, J.M., “Gas conditioning and Processing, Vol 2: The Equipment Modules”, 8th Edition, Edited by R.A. Hubbard, John M. Campbell & Company, Norman, USA, 2000.
    2. ASPENone, Engineering Suite, HYSYS Version 2006, Aspen Technology, Inc., Cambridge, Massachusetts U.S.A., 2006.

  • Acid Gas-Water Phase Behavior

    In the last Tips of the Month, we discussed the phase behavior of water-sweet natural gas and water-sour natural gas mixtures. In this tip, we will demonstrate the acid gas–water phase behavior.

    The water content of a gas depends on the system temperature, pressure and composition of the gas. The phase equilibria in the system H2S + water and CO2 + water are key to the discussion of the water content of an acid gas system. Figure 1 presents the water content of hydrogen sulfide predicted by ProMax [1] as a function of pressure and temperature.  A limited number of experimental data points at 100°F [37.8°C] by Gillespie and Wilson [2] are also shown on this diagram. The behavior shown on this plot is quite complicated and explained by Carroll [3] thoroughly: “At low pressure the hydrogen sulfide + water mixture is in the gas phase. At low pressure the water content tends to decrease with increasing pressure, which is as expected. Eventually a pressure is reached where the H2S is liquefied. On this plot this is represented by the discontinuity in the curve and a broken line joins the phase transition. There is a step change in the water content when there is a transition from vapor to liquid. In the case of hydrogen sulfide the water content of the H2S liquid is greater than the coexisting vapor. This is contrary to the behavior for light hydrocarbons where the water content in the hydrocarbon liquid is less than the coexisting vapor.”

    Graph 1

    Figure 1. Water content of pure H2S predicted by ProMax, experimental data [2]

    In general the phase behavior of the system CO2 + water is as complex as that of the H2S + water system. The CO2-rich liquid phase only occurs for temperatures less than about 90°F [32.2°C]. As shown in Figure 2 reported by Maddox and Lilly [4], the water content of CO2 exhibits a minimum.

    Graph 2

    Figure 2. Predicted saturated water contents at 100°F [38°C] for CO2, CH4 and mixture of both [4]

    There are several methods available that can be used to predict the water content of acid gases. All of these methods are based on equation of state and rigorous thermodynamic models. As described above, the phase behavior is complicated and extra care should be taken to assure a correct prediction. In the remaining section of this tip, we will demonstrate the capabilities of some of these methods.

    Figure 3 compares the water content calculation results for an acid gas stream by several methods in HYSYS [5] and ProMax [1]. The composition of the acid gas stream is shown in the inset of diagram. Even though at low pressures, all methods give close results, as can be seen from this figure, there are large differences at higher pressures.

    Graph 3

    Figure 3. Comparison of water content prediction by different methods at 59 °F [15°C]

    Table 1 gives another comparison of available methods for prediction of acid gas water content.

    Table 1. Comparison of ProMax and modified SRK EOS results with the experimental water content [6] of several acid gas mixtures:

    Table 1

    The experimental composition and predicted water content by HYSYS, ProMax and the modified SRK for eight acid gas mixtures are presented in Table 2. The upper part of this table reports the measured mole percent of the feed stream and the lower part shows the experimental vapor stream compositions in mole percent. Based on the feed compositions, three-phase flash calculations were performed and the resulting vapor stream water content (mole %) are shown in the last three columns (upper part).

    For each vapor stream, the saturated water content was predicted by the above methods and is presented in the lower portion of this table. As can be seen from this table, ProMax predict saturated water content reasonably well. The red figures in Table 2 indicate that the methods predict a non-aqueous liquid phase instead of the vapor phase. Based on a dry basis phase envelope, the conditions for these mixtures were dense phase/compressed liquid.

    Table 2. Conditions and compositions of 8 acid gases and their saturated water contents

    Tables 2 and 3

    Figure 4 also compares the accuracy of the above methods graphically. This figure clearly indicates that ProMax gives the most accurate results. The Erbar et al. [7] SRK method also gives reasonable results.

    Graph 4

    To learn more about similar cases and how to minimize operational problems, we suggest attending

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing – Special), G-6 Gas Treating and Sulfur Recovery and RF-61 Refinery Gas Treating, Sour Water, Sulfur and Tail Gas courses.

    By Wes Wright and M .Moshfeghian

    Reference:

    1. ProMax 2.0, Bryan Research and Engineering, Inc., Bryan, Texas, U.S.A., 2007
    2. Gillespie, P.C. and G.M. Wilson, “Vapor-Liquid Equilibrium Data on Water-Substitute Gas Components: N2-H2O, H2-H2O, CO-H2O, H2-CO-H2O, and H2S-H2O” Research Report RR-41, GPA, Tulsa, OK, 1980.
    3. Carroll, J.J., “The water content of acid gas and sour gas from 100 to 220 °F and pressures to 10,000 Psia,” Presented at the 81st Annual GPA Convention, Dallas, Texas, USA, March 11-13, 2002.
    4. Maddox, R.N., L.L. Lilly, “Gas conditioning and Processing, Vol 3: Computer Applications and Production/Processing Facilities”, John M. Campbell & Company, Norman, USA, 1982.
    5. ASPENone, Engineering Suite, HYSYS Version 2006, Aspen Technology, Inc., Cambridge, Massachusetts U.S.A., 2006.
    6. Huang, S.S.-S., A.-D. Leu, H.-J. Ng, and D.B. Robinson, “The Phase Behavior of Two Mixtures of Methane, Carbon Dioxide, Hydrogen Sulfide, and Water” Fluid Phase Equil. 19, 21-32, 1985.
    7. Erbar, J.H., A.K. Jagota, S. Muthswamy, and M. Moshfeghian, “Predicting Synthetic Gas and Natural Gas Thermodynamic Properties Using a Modified Soave Redlich-Kwong Equation of State,” Gas Processor Research Report, GPA RR-42, Tulsa, USA, 1980.
    8. Ng, H.-J., C.-J. Chen, and H. Schroeder, “Water Content of Natural Gas Systems Containing Acid Gas”,Research Report RR-174, Gas Processors Association, Tulsa, OK, 2001.

  • Water-Sour Natural Gas Phase Behavior

    In the last Tip of the Month, we discussed the phase behavior of water-sweet natural gas mixtures. In this tip, we will demonstrate the water-sour natural gas phase behavior. In a future tip, we will address water content of acid gases.

    Water is produced with oil and gas. A question that comes to mind is: “Why is water important?” The presence of water may cause corrosion, freezing and hydrate formation. All of these problems are enhanced by the presence of acid gases such as H2S and CO2.

    A phase envelope with hydrate and water dew point curves is an excellent tool to find out what form/phase water is in at operating conditions, during start-up, during shut-down and during upsets. The water content of a gas depends on the system temperature, pressure and composition of the water containing gas. There are several methods of calculating of water content of sour gases. The details of these methods can be found in Chapter 6 of Volume 1 [1] and Chapter 9 of Volume 3 [2] of “Gas Conditioning and Processing”. In this work we will use Maddox et al. [3] (Figures 6.1, 2, 3 and Equation 6.2 of Volume 1) and the modified Soave-Redlich-Kwong (SRK EoS) reported in GPA RR-42 by Erbar et al. [4]. This version of SRK is tailor-fitted to handle water-hydrocarbon systems containing hydrogen sulfide and carbon dioxide.

    The compositions of several sour gases studied in this study along with their measured and predicted water contents are shown in Table 1. The Maddox et al. (referred to as Chart) results were generated using GCAP software and the modified SRK EOS results were generated by performing rigorous three-phase (gas-liquid hydrocarbons-aqueous) flash calculations. A trial version of GCAP can be downloaded here, at the bottom of the page.

    Table 1 indicates that as long as the total acid gas concentrations is less than 60 mole percent, Maddox et al. and the modified SRK methods produce results within the accuracy of experimental data. However, for higher concentrations of acid gases, the modified SRK provides a better prediction.  The water content of acid gas systems will be discussed further in the next Tip of the Month.

    Table 1

    The composition, experimental, and predicted water content by Maddox et al. and the modified SRK for two natural gas mixtures are presented in Table 2. The upper part of columns 1, 2, 4, 5 report the measured mole percent. Based on the feed compositions shown in columns 1 and 4, three-phase flash calculations using the modified SRK were performed and the resulting vapor stream compositions are shown in columns 3 and 6, respectively. Notice that the measured and predicted vapor compositions are not identical.  Inaccuracies in predicting the vapor composition can result in errors in predicting the water saturation.

    For each vapor stream, the saturated water content was predicted by both methods and is presented in the lower portion of this table. As can be seen from this table, both methods predict saturated water content reasonably well. Not surprisingly, the accuracy of both methods is improved slightly when the experimental vapor composition is used rather than the predicted vapor composition.

    Table 2

    Figure 1 represents the phase behavior for mixture A. This figure includes from right to left: the water dew point, hydrate formation, 25 weight percent methanol (MeOH) inhibited hydrate formation, hydrocarbon dew point, retrograde, and the bubble point curves, respectively. The blue-triangular-symbol water dew point curve is predicted by use of Figures 6.1, 6.22 and 6.3 with Equation 6.2 of Volume 1 (Maddox et al. method). The red curve represents the water dew point predicted by rigorous calculations using the modified SRK. It is interesting to see that both methods agree quite well with each other. The region to the right of the water dew point curve is gas phase and to the left, liquid water is present.

    Figure 2 presents the phase behavior of sour gas mixture B. With exception of low pressure region, both methods agree quite well.

    Figure 3 demonstrates the effect of acid gases on phase behavior of mixture B. As shown in this figure, the presence of acid gases shifts all of the curves to the right. In other words, the presence of acid gases increases the hydrate formation temperature considerably. It also increases the water dew point temperature. It should be noted that the water dew point curves have been generated for a fixed amount of water content predicted at specified separator condition. In this case the separator temperature and pressure were 120 °F [48.9 °C] and 1500 psia [10,342 kPa], respectively.

    Figure 1

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing – Special) and RF-61 Refinery Gas Treating, Sour Water, Sulfur and Tail Gas courses.

    Dr. Mahmood Moshfeghian

    Reference:

    1. Campbell, J.M., “Gas Conditioning and Processing, Vol 1: The Basic Principles”, 8th Edition, Edited by R.A. Hubbard, John M. Campbell & Company, Norman, USA, 2001.
    2. Maddox, R.N., L.L. Lilly, “Gas conditioning and Processing, Vol 3: Computer Applications and Production/Processing Facilities”, John M. Campbell & Company, Norman, USA, 1982.
    3. Maddox, R.N., L.L. Lilly, M. Moshfeghian, and E. Elizondo, “Estimating Water Content of Sour Natural Gas Mixtures”, Laurance Reid Gas Conditioning Conference, Norman, OK, Mar., 1988.
    4. Erbar, J.H., A.K. Jagota, S. Muthswamy, and M. Moshfeghian, “Predicting Synthetic Gas and Natural Gas Thermodynamic Properties Using a Modified Soave Redlich-Kwong Equation of State,” Gas Processor Research Report, GPA RR-42, Tulsa, USA, 1980.
    5. Huang, S.S.-S., A.-D. Leu, H.-J. Ng, and D.B. Robinson, “The Phase Behavior of Two Mixtures of Methane, Carbon Dioxide, Hydrogen Sulfide, and Water” luid Phase Equil. 19, 21-32, 1985.
    6. Ng, H.-J., C.-J. Chen, and H. Schroeder, “Water Content of Natural Gas Systems Containing Acid Gas”, Research Report RR-174, Gas Processors Association, Tulsa, OK, 2001.

    Figures 2 and 3

  • Water-Sweet Natural Gas Phase Behavior

    In the past Tips of the Month, we discussed the phase behavior of water-free natural gas mixtures. In this tip, we will demonstrate the water-sweet natural gas phase behavior. In future tip, we will address wet sour gas.

    Water is produced with oil and gas. A question that comes to mind is: “why water is important?” The presence of water may cause corrosion, freezing and hydrate formation. Hydrates can even form at warm temperatures in the presence of water. Once hydrates are formed, they are hard to remove. For design and operation of a plant it is important to know:

    1. Where water is in the process?
    2. How much water is present?
    3. What form/phase water is in at operating conditions, during start-up, during shut-down and during upsets?

    A phase envelope with hydrate and water dew point curves is an excellent tool to answer the above questions. The water content of a gas depends on the system temperature, pressure and composition of the water containing gas. There are several methods of calculating of water content. The details of these methods can be found in Chapter 6 of Volume 1 and Chapter 9 of Volume 3 of “Gas Conditioning and Processing” [1, 2]. In this work we will use Figure 6.1 of Volume 1 and the modified Soave-Redlich-Kwong (SRK EoS) reported in GPA RR-42 by Erbar et al. [3]. This version of SRK is tailor fitted for water-hydrocarbon systems. The compositions of natural gases studied in this work are shown in Table 1.

    Table 1

    Figure 1 presents the phase behavior for mixture 1. This figure includes from right to left: the water dew point, hydrocarbon dew point, retrograde, hydrate formation, 25 weight percent methanol (MeOH) inhibited hydrate formation, and the bubble point curves, respectively. The blue-triangular-symbol water dew point curve is predicted by use of Figure 6.1 of Volume 1 and the red curve represents the water dew point predicted by rigorous calculations using the modified SRK. It is interesting to see that both methods agree quite well with each other. However, the results obtained by Figure 6.1 are somewhat more conservative. The region to the right of water dew point curve is gas phase and to the left the liquid water is present.

    Graphs 1 and 2

    Figure 2 presents the phase behavior for mixture 2 which contains heavier compounds including nC9H20. However, the water content is the same as in mixture 1. Note the position of the water dew points did not change but, the hydrocarbon dew point curve has moved to the right, as expected, due to presence of heavier compounds. At lower pressures, the water dew point curves coincide with the hydrocarbon dew point curve. This is merely by coincidence.

    Figure 3 presents similar phase diagram for mixture 3 which is essentially the same as mixture 1 except nC6H14 has been replaced with nC9H20. Again, the hydrocarbon dew point curve shifts to the right due to presence of nC9H20.

    Graph 3

    It is interesting to note that the dry hydrocarbon dew points and the wet hydrocarbon dew points predicted by SRK coincide very closely with each other; the difference is practically negligible. Also note that at a specified pressure, the higher of the two dew points (hydrocarbon and water) have been calculated by the SRK EoS. So, below about 1400 psia [9653 kPa], the wet hydrocarbon dew point is predicted while for pressures above 1400 psia [9653 kPa], the water dew point (the higher one) is predicted.

    Finally, mixture 2 has been passed through a separator at 100 °F [38 °C] and 1000 psia [6895 kPa] and the resulting vapor compositions from a three-phase flash calculation based on the SRK EoS is shown in the last column of Table 1 as mixture 4. Due to the removal of free water and heavy hydrocarbons from mixture 2, the phase envelope and the water dew point curve have moved to the left, as expected. At this condition, the water content by SRK EoS is 0.0012 mole fraction equivalent of 57 lbm/MMSCF or 914 kg/106 std. m3. Figure 4 indicates that the hydrocarbon dew point and water dew point curves intersect at 100 °F [38 °C] and 1000 psia [6895 kPa] which are the conditions of the separator.

    Due to the fact that hydrate formation is controlled mostly by lighter components, there are only small variations of the hydrate formation curve and its inhibition by 25 weight percent methanol in all four mixtures.

    To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By: Dr. Mahmood Moshfeghian

    Graph 4

    Reference:

    1. Campbell, J.M., “Gas conditioning and Processing, Vol 1: The Basic Principles”, 8th Edition, Edited by R.A. Hubbard, John M. Campbell & Company, Norman, USA, 2001.
    2. Maddox, R.N., “Gas conditioning and Processing, Vol 3: Computer Applications and Production/Processing Facilities”, John M. Campbell & Company, Norman, USA, 1982.
    3. Erbar, J.H., Jagota, A.K., Muthswamy, S. and Moshfeghian, M., “Predicting Synthetic Gas and Natural Gas Thermodynamic Properties Using a Modified Soave Redlich Kwong Equation of State,” Gas Processor Research Report, GPA RR-42, Tulsa, USA, 1980.

  • Consequence of Liquid Carry Over – Part 2: Fixed Heat Exchanger Area

    Many facility operating problems are related to improperly designed or under-sized gas-liquid separators. Due to the importance of separators, in the July Tip of the Month (TOTM), we studied the effect of liquid carry over in a simple dew point control plant. In that study, assuming variable area of the heat exchangers, we found that one percent liquid carry over can cause considerable change in compressor power requirement and heat exchanger duty. This assumption is valid if we are designing a new plant and the equipment is being sized for construction. However, for an existing plant, the heat exchanger areas are fixed. In the continuation of our previous study, we will revisit the same case study and investigate the consequence of liquid carry over but we will use the same heat exchangers (i.e. keeping UA constant; U=overall heat transfer coefficient and A=area).

    Let’s consider the same process flow diagram as shown in Figure 1 for a simple gas plant. The feed composition and conditions are shown in Table 1. We have already shown the impact of liquid carry over on the “sales gas dew point” and we experienced that for even a small liquid carry over (1 %), the dew point offset was about 6°F [3.3 °C]. We will demonstrate the consequences of liquid carry over on the other equipment while maintaining the “spec dew point”.

    Process Flow Diagram

    Figure 1. Process flow diagram for simple dew point correction gas plant

    It is desired to process this feed gas to produce a sales gas with a dew point of 20°F [-6.7 °C] at 540 psig [3.723 MPa] The feed gas is mixed with recycle gas from a stabilizer, compressed and cooled to 110°F [43.3°C] and 555 psig [3.827 MPa], then cooled in the gas-gas exchanger, gas-liquid exchanger and finally in the chiller to 20°F [-6.7°C] before entering the separator at 540 psig [3.723 MPa]. In real equipment, there would be some liquid carry over. In order to show the impact of liquid carry over, in the simulation we withdraw a small portion of liquid stream from the separator and remixed it with the vapor stream.  The solid curve in Figure 2 shows how the dew point of sales gas goes off spec as a function of the liquid carry over.

    In order to bring back the sales gas dew point to spec, we re-adjusted (lowered) the stream 7 temperature. The required degree of re-adjustment is shown by the dashed line in the same figure. As a consequence of re-adjusting of the stream 7 temperature, the conditions of other equipment and streams changed. Figure 3 shows the variation of compressor power and the heat exchanger duties as a function of liquid carry over. As can be seen in this figure, one percent liquid carry over can cause considerable change. For this case, the percent changes ranged from 0 to 27 percent. The changes in the reboiler duty, sales gas and LPG flow rates were negligible. Contrary to the findings of the July TOTM, here we see that the chiller duty increases as we expected.

    Not included in this analysis is an examination of the affect of lowering the chiller outlet temperature on the refrigeration system.  In an existing plant, to lower the refrigerant temperature, the chiller would have to operate at a lower pressure so that the power required for the refrigerant compressor would increase. The overall effect of liquid carry over is the increase in the operating cost, as expected.

    To learn more about similar cases and how to minimize operational problems such as liquid carry over, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By: Dr. Mahmood Moshfeghian

    Tables 1 and 2Graphs 1 and 2

  • Consequence of Liquid Carry Over in a Simple Dew Point Control Plant

    Problems in meeting sales-gas dew point specifications are not unusual. A facility engineer often suspects separator carryover when trouble-shooting such a plant. Proper sizing of equipment for vapor-liquid separation is essential to almost all processes. The fundamentals of a simple separator design may be extended to several other processes such as fractionation towers, two-phase flow, slug catcher design etc. Many facility operating problems are related to improperly designed or under-sized gas-liquid separators.

    Let’s consider the process flow diagram shown in Figure 1 for a simple gas plant. The feed composition and conditions are shown in Table 1. In one of our past Tip of the Months (TOTM) we demonstrated the impact of liquid carry over on the “sales gas dew point” and we experienced that for even a small liquid carry over (1 %), the dew point offset was about 6°F [3.3 °C]. In this Tip of the Month, we will demonstrate the consequences of liquid carry over on the other equipment while maintaining the “spec dew point”.

    Process Flow Diagram

    It is desired to process this feed gas to produce a sales gas with a dew point of 20 °F (-6.7 °C) at 540 psig (3.723 MPa) The feed gas is mixed with recycle gas from a stabilizer, compressed and cooled to 110°F (43.3°C) and 555 psig (3.827 MPa), then cooled in the gas-gas exchanger, gas-liquid exchanger and finally in the chiller to 20°F (-6.7°C) before entering the separator at 540 psig (3.723 MPa). In real equipment, there would be some liquid carry over. The commercial simulators will assume a perfect gas-liquid separation unless the users manually force some carryover. In order to show the impact of liquid carry over, in simulation we withdraw a small portion of liquid stream from separator and remixed it with the vapor stream. The solid curve in Figure 2 shows how the dew point of sales gas goes off spec as a function of the liquid carry over.

    In order to bring back the sales gas dew point to spec, we re-adjusted (lowered) the stream 7 temperature. The required degree of re-adjustment is shown by the dashed line in the same figure. As a consequence of re-adjusting of the stream 7 temperature, the conditions of other equipment and streams changed. Figure 3 shows the variation of compressor power and the heat exchanger duties as a function of liquid carry over. As can be seen in this figure, a-one percent liquid carry over can cause considerable change. For this case, the percent changes ranged from 15 to 55 percent. The changes in the reboiler duty, sales gas and LPG flow rates were negligible. Please note that we assumed variable area of the heat exchangers. If this analysis is done prior to building a new plant, the largest heat exchangers needed could be purchased. However, in an existing plant the heat exchanger areas are fixed. An interesting “surprise” can be seen in Figure 3, the duty of chiller went down as the liquid carryover increased. This is due to the fact that the enthalpy of stream 6 decreases more than the required change in the enthalpy (temperature) of the cold separator. Therefore the process gas duty across the chiller decreases as the liquid carry over increases (see Table 2). In other words, liquid carry-over from the LTS makes the gas/gas heat exchanger into a “chiller” as the liquids vaporizes in that heat exchanger, lowering actual chiller duty, but still increasing the sales gas dewpoint temperature.

    Not included in this analysis is an examination of the affect of lowering the chiller outlet temperature on the refrigeration system. In an existing plant, to lower the refrigerant temperature, the chiller would have to operate at a lower pressure so that the power required for the refrigerant compressor would increase. The overall effect of liquid carry over is the increase in the operating cost, as expected.

    To learn more about similar cases and how to minimize operational problems such as liquid carry over, we suggest attending our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By: Dr. Mahmood Moshfeghian

    Tables 1 and 2 

    Graphs 1 and 2

  • MEG Dehydration Ability in MEG Injection Plant

    In order to continue the last tip of the month’s discussion on MEG injection plant, in this “Tip of the Month”, we will focus on two more questions:

    1. Does MEG have any dehydration ability at the three phase cold separator condition of a typical mechanical refrigeration plant?
    2. What is the dehydration ability of MEG if the mechanical refrigeration in a typical MEG injection plant goes out of service?

    As described in the last tip of the month, a typical mechanical refrigeration process is used for hydrocarbon dew point control and moderate NGL recovery that uses MEG injection to prevent hydrate formation. Warm inlet gas is cross-exchanged with the cold dry sales gas and then flows to the gas chiller. To prevent hydrates from forming, MEG is injected in the tubes at the warm end of both exchangers. The temperature of the chiller is adjusted to condense liquids from the feed gas. The cold gas exiting the chiller together with the rich MEG solution and condensed hydrocarbons enters the cold three-phase separator. The rich MEG is sent to the regeneration section of the unit where the water is removed. The resulting lean MEG is sent back to the process. In short, two things are taking place: temperature reduction of the process gas to condense both water and hydrocarbons; and, MEG injection and subsequent regeneration to prevent hydrates from forming. In this scheme, the sales gas exiting the gas-to-gas exchanger has a water and hydrocarbon dew point determined by the operating temperature of the cold separator. The key point to remember here is that the water is being removed from the gas by low temperature condensation. The purpose of the injected MEG is not to “dehydrate” the gas but to prevent formation of hydrates. For more detail, refer to chapters 6 and 16 of Gas Conditioning and Processing, Volumes 1 and 2, [1, 2] respectively.

    Question 1: Does MEG have any dehydration ability at the three phase cold separator condition of a typical mechanical refrigeration plant?

    In order to answer this question, first we determine the water dew point temperature without any MEG injection and compare the results with the case of 80 weight percent lean MEG injection. Let’s assume a typical natural gas, cold separator pressure of 40 bara and -20°C [580 psia & -4°F] with 10 weight dilution (i.e. rich MEG concentration of 70 weight %). By performing computer simulation using ProMax [3], the water dew point temperature:

    • without MEG injection is -22.7°C (-8.7°F) corresponding to water content of 30.72 kg/106 std m3 [1.94 lbm/MMSCF]
    • with MEG injection is -29.6°C (-21.2°F) corresponding to water content of 17.6 kg/106 std m3 (1.11 lbm/MMSCF)

    Therefore, the water dew point temperature depression is 6.9°C (12.4°F). Similarly, a water dewpoint temperature depression of 7.8°C [14°F] was obtained for the case of 5 weight percent MEG dilution (i.e. rich MEG concentration of 75 weight %). These results indicate that even at low temperature, in addition to the hydrate inhibition effect, MEG has the ability to do partial dehydration. It should be noted that for this gas the hydrate formation temperature at 40 bara [580 psia] is 14.9°C [58.7°F].

    Question 2: What is the dehydration ability of MEG if the mechanical refrigeration in a typical MEG injection plant unexpectedly goes out of service?

    Let’s assume the same gas as in question 1 is passing through a mechanical refrigeration system with the same chiller temperature of -20°C [-4°F]. Let’s also assume that due to the break down of mechanical refrigeration system (lack of chilling) the cold separator temperature reaches 21.1°C [70°F]. Again, we used ProMax to perform the simulations and the calculated results are plotted in Figures 1 (A&B) and 2 (A&B). Figure 1 (A&B) presents the effect of lean MEG circulation rate on water dew point temperature and water content. Figure 2 (A&B) indicates that for a 10 weight % dilution, about 4350 kg MEG solution per 106 std m3 [270 lbm MEG solution per MMSCF] of gas is required. Figure 2 (A&B) also indicates that for this amount of dilution, the water dew point temperature drops from 21.1°C [70°F] to about 12.2°C [54°F] and the corresponding water content drops from 567 to 325 kg/106 std m3 [35 to 21 lbm/MMSCF]. Again, it can be seen that the MEG can dehydrate natural gas partially at higher temperature. It is also interesting to see from Figures 1 and 2 that further increase in lean MEG solution circulation rate, beyond 4350 kg/106 std m3 [270 lbm/MMSCF], does not reduce the water dewpoint temperature considerably and; therefore, it justifies the rule of thumb for 10 weight % dilution.

    For more information about dehydration and hydrate inhibition, the reader should refer to JMC books and enroll in our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By Dr. Mahmood Moshfeghian

    References:

    1. Campbell, J. M. “Gas conditioning and processing, Volume 1: Basic Principles,” 8th Ed., John M. Campbell and Company, Norman, Oklahoma, USA, 2001.
    2. Campbell, J. M. “Gas conditioning and processing, Volume 2: The Equipment Modules,” 8th Ed., John M. Campbell and Company, Norman, Oklahoma, USA, 2000.
    3. ProMax, version 1.2, Bryan Research & Engineering Inc, Bryan, Texas, 2005.

    Figure 1AFigure 1BFigure 2AFigure 2B

  • MEG Injection vs. TEG Dehydration

    In this “Tip of the Month”, we will focus on the question of: Which technology should you choose? The answer, of course, is “It depends.” It depends on what you are trying to accomplish, the constraints imposed on your system and the relative economics.

    A Rule of Thumb is “Use MEG injection if you have to cool the gas for NGL recovery anyway.” Like all Rules of Thumb, there are exceptions. But let’s explore the basics of each technology.

    Let’s begin by defining our terms. See Figure 1 for a typical mechanical refrigeration process used for hydrocarbon dew point control and moderate NGL recovery that uses MEG injection to prevent hydrate formation. Warm inlet gas is cross-exchanged with the cold dry sales gas and then flows to the gas chiller. To prevent hydrates from forming, MEG is injected in the tubes at the warm end of both exchangers. The temperature of the chiller is adjusted to condense liquids from the feed gas. The cold gas exiting the chiller together with the rich MEG solution and condensed hydrocarbons enters the cold three-phase separator. The rich MEG is sent to the regeneration section of the unit where the water is removed. The resulting lean MEG is sent back to the process.

    Process Flow Diagram

    Copyright © 2007 John M. Campbell and Company

    Figure 1. Typical mechanical refrigeration plant with glycol injection system [1]

    In this flow diagram, two things are taking place: temperature reduction of the process gas to condense both water and hydrocarbons; and, MEG injection and subsequent regeneration to prevent hydrates from forming. Inspection of Figure 1 reveals the majority of the equipment, including the refrigeration compressors, etc. which are not shown, is employed to reduce the temperature. Besides mechanical refrigeration, other options to achieve the required gas cooling include JT – valve expansion or use of a turboexpander. For either of these options, the MEG injection and regeneration portion of this plant is minor by comparison.

    In this scheme, the sales gas exiting the gas-to-gas exchanger has a water and hydrocarbon dew point determined by the operating temperature of the cold separator. The CAPEX of this system is essentially driven by the gas cooling equipment, including the refrigeration system. The key point to remember here is that the water is being removed from the gas by low temperature condensation. The purpose of the injected MEG is not to “dehydrate” the gas but to prevent formation of hydrates. At the MEG concentrations normally used in these systems, approximately 80 – 85 wt%, the MEG absorbs only a small amount of water vapor from the gas.

    Let’s now look at a typical circulating TEG system. See Figure 2. The same rich, water saturated natural gas stream flows to a properly sized inlet separator to remove liquids. The gas then enters a glycol contactor equipped with either structured packing or bubble cap trays. As the gas rises, the water is removed by the falling TEG. The concentration of the lean glycol entering the top of the contactor is the main variable that determines the water dew point specification that can be made. The rich glycol that leaves the glycol contactor is sent to a flash drum and then to a regeneration section. The lean glycol leaving the regenerator is then returned to the contacting tower.

    In this system, we are only making water dew point specification gas. The NGL content/hydrocarbon dew point of the sales gas is the same as that of the feed gas. Circulating TEG systems are therefore used only for dehydration. A significant cost item for the circulating TEG system is the high pressure contacting tower.

    Now let’s explore how we can compare and contrast these two technologies.

    If your objective is to make only pipeline water specification gas, you will most likely choose a circulating TEG system. This is intuitively obvious from a comparison of the two flow diagrams cited above. Assume, for example, that you want to dehydrate a lean natural gas stream that is water saturated at 70 bar and 40°C. A quick comparison of Figures 1 and 2 shows that there is much more equipment associated with chilling the feed gas (Figure 1 + the refrigeration compressors, etc. that are not shown) then there is with a circulating TEG system (Figure 2). Hence, for dehydration only to pipeline water specifications, a circulating TEG system will almost always be selected.

    On the other hand, if your objective is to recover hydrocarbons and remove water simultaneously, then a low – temperature process with MEG injection may be the best choice. Assume you have a rich natural gas stream that is water saturated at 70 bar and 40°C. Assume a mechanical refrigeration process is selected for hydrocarbon liquids recovery with a cold temperature of -35°C. We have two options to consider: we can dehydrate the gas with a circulating TEG system to a water dew point of -35°C and then send the dehydrated gas to an LTS plant consisting of a gas-to-gas exchanger, chiller, refrigeration system, etc., but with no MEG injection/regeneration system; or, we can send the feed gas directly to the LTS plant which has an MEG injection system retrofitted to prevent hydrates from forming.

    Figure 2

    Copyright © 2007 John M. Campbell and Company

    Figure 2. Basic glycol dehydration unit [2]

    Since the underlying equipment required to recover NGL’s is the same in both options, the cost comparison is essentially between the circulating TEG system and the MEG injection system. The TEG system will use less circulating rates then the MEG system, but will likely have a higher regeneration duty. Achieving the large dew point depression of 75°C with a circulating TEG system will be challenge and will add to the system cost. The key difference, however, is the circulating TEG system requires a high pressure contactor while the MEG injection system does not. In this situation, the most likely choice will be to go with the MEG Injection system.

    For more information about dehydration and hydrate inhibition, the reader should refer to JMC books and enroll in our G4 (Gas Conditioning and Processing) and G5 (Gas Conditioning and Processing – Special) courses.

    By Harvey M. Malino and Mark Bothamley

    References:

    1. Campbell, J. M. “Gas conditioning and processing, Volume 1: Basic Principles,” 8th Ed., John M. Campbell and Company, Norman, Oklahoma, USA, 2001.
    2. Campbell, J. M. “Gas conditioning and processing, Volume 2: The Equipment Modules,” 8th Ed., John M. Campbell and Company, Norman, Oklahoma, USA, 2000.