Campbell Tip of the Month

Author: Dr. Mahmood Moshfeghian

Variation of Crude Oil Properties with Temperature in a Gathering Line
In the October and November 2014 Tips of the Month (TOTM), we demonstrated that Gas-Oil-Ratio (GOR) has a large impact on the capacity of crude oil gathering lines. If GOR is less than the saturation solution gas, the increase in GOR reduces the viscosity and density of crude oil which causes the pressure drop to decrease. However, if the GOR exceeds the saturation solution gas the system becomes two phase and pressure drop increases.

The solution gas is a function of temperature, pressure, gas and liquid compositions. In this TOTM, we will study the impact of temperature on the crude oil properties in the gathering systems for the case presented in the November 2014 TOTM. Specifically, the variation of the crude oil relative density and viscosity with the temperature will be studied. Finally, the impact of temperature on the oil and gas velocity and pressure drop along a gathering line for nominal pressure of 6900 kPag (1000 psig) and nominal pipe size of 101.6 mm (4 inches) will be demonstrated using a multiphase rigorous method from a commercial simulator. The calculated properties, oil and gas velocities and pressure drops are presented in graphical format as a function of the oil stock tank volume flow rate, solution gas, R_S, and temperature.

Case Study
1. For the purpose of illustration, we considered a case study for transporting a crude oil of relative density of 0.852 (°API = 34.6) at stock tank conditions combined with a gas with relative density of 0.751. The selected GORs were 0 (dead oil), 17.8, 35.6, and 89 Sm³ of gas/STm³ of oil (0, 100, 200, 500 scf/STB). The compositions of oil and gas are presented in Table 1. The oil C₆₊ was characterized as 30 hypothetical single carbon number (SCN) [1] ranging from SCN6 to SCN35 while the gas C₆₊ was characterized by 10 hypothetical SCN ranging from SCN6 to SCN15. For details of the SCN components, see Table 3.2 on page 64 of reference [1]. The mole fraction of SCN components were determined by an exponential decay algorithm [2]. The feed enters the line at 15.6 ˚C (60 ˚F) for case 1 and 43.3 ˚C (110 ˚F) for case 2.
  Table 1. Feed composition at stock condition
The following assumptions were made:
1. Steady state conditions
2. The line is 1.601 km (1 mile) long with nominal size of 101.6 (4 inches), onshore buried line.
3. Segment lengths and elevation changes are presented in Table 2 and Figure 1. This elevation profile is considered to be approximately equivalent to “rolling” terrain.
4. Pipeline inside surface roughness of 46 microns (0.046 mm, 0.0018 inch)
5. Line nominal pressure 6900 kPag (1000 psig)
6. The feed enters the line at 15.6 ˚C (60 ˚F) for case 1 and 43.3 ˚C (110 ˚F) for case 2.
7. The ground/ambient temperature, is 15.6 ˚C (60 ˚F)
8. Water cut is 0 (no water in the feed).
9. Overall heat transfer coefficients of 2.839 W/m²-˚C (0.5 Btu/hr-ft²-˚F), for onshore buried line.
10. Simulation software ProMax [3] and using the Soave-Redlich-Kwong (SRK) Equation of State [4] for vapor-liquid equilibrium and Beggs-Brill method for two-phase pressure drop calculation [5].
Table 2. Line segment length and elevation change

Figure 1. Gathering line elevation profile

Results and Discussions:

The two phase (oil and gas) flow through the gathering line was simulated by ProMax with SRK EOS for vapor-liquid equilibria and Beggs-Brill for two phase pressure drop calculations. Figure 2 presents the variation of temperature along the pipeline for the case of 0 GOR for an oil rate of 636 STm³/d (4000 STB/day). Since the ambient temperature was assumed to be 15.6°C (60°F), for case 1 the crude oil temperature in the line remained constant. However, for the second case the temperature dropped from 43.3°C (110°F) to about 40°C (104°F) within 1.6 km (1 mile) of the line. For the second case the average line temperature is about 41.7°C (107°F).

Figure 2. Variation of line temperature along gathering line for two feed temperatures of 15.6 and 43.3 C (60 and 110) and an oil rate of 636 STm³/d (4000 STB/day)

Figure 3 presents the bubble point pressure of the feed to the gathering line at the average line temperatures of 15.6 and 41.7 (60 and 107) as a function of solution gas. This figure indicates that for the nominal line pressure of 6900 kPa (1000 psig), the crude oil is under saturated up to GOR of 51.8 Sm³/STm³ (290.5 scf/STB) for the lower average line temperature. Similarly it shows oil is under saturated up to GOR of 39.4 Sm³/STm³ (221.3 scf/STB) for the higher average line temperature. For GOR greater than these values, the oil becomes saturated with gas and gas breaks out. The system becomes two phase gas and liquid flow.

Figure 4 presents the line pressure drop per unit length as a function of oil stock tank volume rate, GOR, and feed temperature. In this figure and subsequent figures, the solid lines are for the feed temperature of 15.6°C (60°F) and symbols are for feed temperature of 43.3°C (107°F). Figure 4 indicates that as the GOR increases from 0 to 35.7 Sm³/STm³ (0 to 200 scf/STB), the pressure drop decreases but increases with further increase in GOR of 89 Sm³/STm³ (500 scf/STB) and higher. The dividing point is at a saturation solution gas of 39.4 Sm³/STm³ (221.3 scf/STB) and 51.8 Sm³/STm³ (290.5 scf/STB) for temperatures of 41.7°C (104°F) and 15.6°C (60°F), respectively. At higher temperature the increase of GOR reduces the pressure drop when solution gas is under saturated but increases the pressure drop for GOR greater than the saturated solution gas.

Figure 3. Bubble point pressure of the feed to the gathering line as a function of solution gas at nominal line temperature of 15.6 and 41.7 (60 and 107)

Figure 4. Variation of pressure drop per unit length with oil stock tank volume rate, GOR, and temperature – Solid curves for 15.6°C (60°F) and symbols for 43.3°C (107°F)

Figure 5 presents the variation of oil relative density along the line as a function of solution gas (R_S) and temperature. This figure indicates that as the R_S increases, the oil relative density decreases. Note as the temperature increases the solution gas (R_S) decreases.

Figure 5. Variation of oil relative density with solution gas and temperature along the gathering line at 6900 kPag for 101.6 mm pipe diameter, oil rate of 636 STm³/d (4000 STB/day). Lines are for feed temperatures of 15.6 and symbols are for 43.3 C (60 and 110)

Figure 6 shows that as the R_S increases, the oil viscosity decreases considerably. The reduction of viscosity causes pressure drop to decrease. The simulation results (Figure 3) indicated that for GOR less than 51.8 Sm³/STm³ (290.5 scf/STB) at 15.6°C (60°F) or 39.4 Sm³/STm³ (221.3 scf/STB) at 41.7°C (104°F), the flow is under saturated single liquid phase; however, for higher GOR the flow becomes saturated two phase (gas and liquid) which causes the pressure drop to increase. The increase in pressure drop due to higher GOR (and higher total flow rate) is more than the decrease in pressure drop due to reduction of oil viscosity as a result of solution gas. The net effect is higher pressure drop compared to dead oil (GOR = 0) pressure drop. This figure also indicates that at zero and low GOR, in which the system is single liquid phase, the temperature has a larger impact on the crude oil viscosity.

Figure 6. Variation of oil viscosity with solution gas, along the gathering line at 6900 kPag for 101.6 mm pipe diameter, oil rate of 636 STm³/d (4000 STB/day). Lines are for feed temperatures of 15.6 and symbols are for 43.3 (60 and 110)

Figure 7 presents the variation of oil and gas velocity for two stock tank oil rates and two feed temperatures along the gathering line at 6900 kPag for 101.6 mm pipe diameter and GOR of 89.1 Sm³/STm³ (500 scf/STB). The lines are for temperature of 15.6 and symbols are for temperature of 43.3 C (60 and 110). Figure 7 indicates that the oil velocity remains constant along the line but the gas velocity increases due to pressure drop in the line and more gas coming out of the solution. This figure also indicates that the impact of temperature on gas velocity is more than on the liquid phase velocity.

Figure 8 presents the impact of GOR and temperature on pressure drop along the gathering line at 6900 kPag for 101.6 mm pipe diameter and an oil rate of 636 STm³/d (4000 STB/day).

Figure 7. Variation of oil and gas velocity for two stock tank oil rate along the gathering line at 6900 kPag for 101.6 mm pipe diameter and GOR of 89 Sm³/STm³ (500 scf/STB). Lines are for feed temperatures of 15.6 and symbols are for 43.3 C (60 and 110)

Figure 8. Impact of GOR, Sm³/STm³ (scf/STB) on pressure drop along the gathering line at 6900 kPag for 101.6 mm pipe diameter, oil rate of 636 STm³/d (4000 STB/day). Lines are for feed temperatures of 15.6 and symbols are for 43.3 C (60 and 110)

As can be seen in this figure, for GOR less than saturation GOR, the pressure drop decreases as GOR increases but for GOR greater than saturation GOR, due to presence of two phase flow, the pressure drop increases. It also shows that for GOR less than saturation, increase in temperature reduces the pressure drop but for GOR greater than saturation, the increase in temperature increases the pressure drop.

Conclusions

The following conclusions can be made based on this case study:
1. For under saturated oil, the increase in temperature reduces oil density, viscosity and pressure drop. For saturated oil, the increase in temperature increases pressure drop due to increased free gas volumes.
2. Even though increasing temperature reduces the solution gas, it reduces the viscosity of under saturated crude oil but its impact on saturated liquid viscosity diminishes. Increasing temperature decreases the viscosity of crude oil but it also decreases the gas in solution. The two effects can diminish the effect of temperature on viscosity.
3. The impact of temperature on gas velocity is higher than its impact on liquid velocity.
To learn more about similar cases and how to minimize operational problems, we suggest attending our PF 45 (Onshore Gas Gathering Systems: Design and Operation), G4 (Gas Conditioning and Processing), PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Mahmood Moshfeghian

Reference:
1. Campbell, J.M., Gas Conditioning and Processing, Volume 1: The Basic Principles, 9^th Edition, 2^nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
2. Moshfeghian, M., Maddox, R.N., and A.H. Johannes, “Application of Exponential Decay Distribution of C₆₊ Cut for Lean Natural Gas Phase Envelope,” J. of Chem. Engr. Japan, Vol 39, No 4, pp.375-382 (2006)
3. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
4. Soave, G., Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.
5. Brill, J. P., et al., “Analysis of Two-Phase Tests in Large-Diameter Flow Lines in Prudhoe Bay Field,” SPE Jour, 363-78, June 1981.
January 1, 2015
Impact of Solution Gas on Crude Oil Properties in a Gathering Line
In the October 2014 Tip of the Month (TOTM), we demonstrated that Gas-Oil-Ratio (GOR) has a large impact on the capacity of crude oil gathering lines. In general as GOR increased the pressure drop increased which lowered the line capacity. In addition, at high pressures and low GOR, pressure drop was lower than the pressure drop for dead oil (solution gas is zero) because the viscosity of live oil is lower than viscosity of dead oil. This effect was bigger for the smaller line diameter.

In this TOTM, we will study the impact of solution gas (RS) on the crude oil properties in the gathering systems for one of the cases presented in the October 2014 TOTM. Specifically, the variation of the crude oil relative density and viscosity with the solution gas (RS) will be studied. Finally, the impact of solution gas (RS) on the oil and gas velocity and pressure drop along a gathering line for nominal pressure of 6900 kPag (1000 psig) and nominal pipe size of 101.6 mm (4 inches) will be demonstrated using multiphase rigorous method from a commercial simulator. The calculated properties, oil and gas velocities and pressure drops are presented in graphical format as a function of the oil stock tank volume flow rate and solution gas, RS.

For clarity, gas-oil-ratio (GOR) is defined as the total volume of gas which comes out of oil at standard conditions divided by the total volume oil at the stock tank conditions. The solution gas (RS) is defined as the volume of gas dissolved in oil divided by the volume of oil (with the same unis as GOR) but at the flowing temperature and pressure.

Case Study

For the purpose of illustration, we considered a case study for transporting a crude oil of relative density of 0.852 (°API = 34.6) at stock tank conditions combined with a gas with relative density of 0.751. The selected GORs were 0 (dead oil), 17.8, 35.6, and 89 Sm3 of gas/STm3 of oil (0, 100, 200, 500 scf/STB). The compositions of oil and gas are presented in Table 1. The oil C6+ was characterized as 30 hypothetical single carbon number (SCN) [1] ranging from SCN6 to SCN35 while the gas C6+ was characterized by 10 hypothetical SCN ranging from SCN6 to SCN15. For details of the SCN components, see Table 3.2 on page 64 of reference [1]. The mole fraction of SCN components were determined by an exponential decay algorithm [2].

Table 1. Feed composition at stock condition
The following assumptions were made:

Steady state conditions

The line is 1.601 km (1 mile) long with nominal size of 101.6 (4 inches), onshore buried line.

Segment lengths and elevation changes are presented in Table 2 and Figure 1. This elevation profile is considered to be approximately equivalent to “rolling” terrain.

Pipeline inside surface roughness of 46 microns (0.046 mm, 0.0018 inch)

Line nominal pressure 6900 kPag (1000 psig)

The feed enters the line at 15.6 ̊C (60 ̊F)

The ground/ambient temperature, is 15.6 ̊C (60 ̊F)

Water cut is 0 (no water in the feed).

Overall heat transfer coefficients of 2.839 W/m2- ̊C (0.5 Btu/hr-ft2- ̊F), for onshore

buried line (minor effect as inlet temperature = ambient ground temperature).

Simulation software ProMax [3] and using the Soave-Redlich-Kwong (SRK) Equation of State [4] for vapor-liquid equilibrium and Beggs-Brill method for two-phase pressure drop calculation [5].
Table 2. Line segment length and elevation change
Figure 1. Gathering line elevation profile

Results and Discussions:

The two phase (oil and gas) flow through the gathering line was simulated by ProMax with SRK EOS for vapor-liquid equilibria and Beggs-Brill for two phase pressure drop calculations. Figures 2A and 2B present the calculated pressure drop per unit length as a function of oil stock tank volume rate and GOR for nominal line diameter of 101.6 mm (4 inches) at nominal line pressure of 6900 kPag (1000 psig) in SI (international) and FPS (Engineering) system of units, respectively. Figures 2A and 2B indicate that as the GOR increases from 0 to 35.7 Sm³/STm³ (0 to 200 scf/STB), the pressure drop decreases but increases with further increase in GOR of 89 Sm³/STm³ (500 scf/STB) and higher.

The impact of R_S on the properties of oil is demonstrated in the next section which will explain the causes of pressure drop.

Figure 2A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 6900 kPag for 101.6 mm pipe diameter

Figure 2B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 1000 psig for 4 in pipe diameter

Figure 3 presents the bubble point pressure of the feed to the gathering line at 15.6 (60) as a function of solution gas. Figure 3 indicates that for the nominal line pressure of 6900 kPa (1000 psig), the feed is under saturated up to GOR of 51.8 Sm³/STm³ (290.5 scf/STB). For GOR greater than this value, the oil becomes saturated with gas and gas breaks out.

Figure 3. Bubble point pressure of the feed to the gathering line as a function of solution gas at 15.6 (60)

The variation of oil relative density along the line as a function of solution gas (R_S), is presented in Figure 4. This figure indicates that as the R_S increases, the oil relative density decreases. Figure 5 shows that as the R_S increases, the oil viscosity decreases considerably. The reduction of viscosity causes pressure drop to decrease. The simulation results (Figure 3) indicated that for GOR less than 51.8 Sm³/STm³ (290.5 scf/STB), the flow is under saturated single liquid phase; however, for higher GOR the flow becomes saturated two phase (gas and liquid) which causes the pressure drop to increase. The increase in pressure drop due to higher GOR (and higher total flow rate) is more than the decrease in pressure drop due to reduction of oil viscosity as a result of solution gas. The net effect is higher pressure drop compared to dead oil (GOR = 0) pressure drop.

Figure 6 presents the variation of oil and gas velocity for two stock tank oil rate along the gathering line at 6900 kPag for 101.6 mm pipe diameter and GOR of 89.1 Sm³/STm³ (500 scf/STB). Figure 6 indicates that the oil velocity remains constant along the line but the gas velocity increases due to pressure drop in the line.

Figure 7 presents the impact of GOR on pressure drop along the gathering line at 6900 kPag for 101.6 mm pipe diameter. As it can be seen in this figure, for GOR less than 51.8 Sm³/STm³ (290.5 scf/STB) the pressure drop decreases as GOR increases but at higher GOR due to presence of two phase flow, the pressure drop increases. As the GOR increases further, the effect of elevation change diminishes compared to rise of pressure drop due to friction.

Conclusions

The following conclusions can be made based on this case study:
1. If the oil is under saturated the increase in solution gas (R_S), reduces the oil viscosity and causes the pressure drop to decrease. For saturated oil, the increase in GOR changes the single phase liquid flow to two phase gas-liquid flow and causes the pressure drop to increase and overcome the pressure drop reduction due to lower liquid viscosity.
2. While increasing solution gas (R_S) reduces the oil viscosity and relative density they remain almost constant along the line.
3. While at higher GOR the flow becomes two phase, the pressure drop due to friction becomes dominant and overcomes the elevation changes. This is more pronounced in the longer lines.
4. While oil velocity remains constant along the line, the gas velocity increases along the line.
Figure 4. Variation of oil relative density with solution gas (R_S), Sm³/STm³ (scf/STB), along the gathering line at 6900 kPag for 101.6 mm pipe diameter

Figure 5. Variation of oil viscosity with solution gas (R_S), Sm³/STm³ (scf/STB), along the gathering line at 6900 kPag for 101.6 mm pipe diameter

To learn more about similar cases and how to minimize operational problems, we suggest attending our PF45 (Onshore Gas Gathering Systems: Design and Operation), G4 (Gas Conditioning and Processing), PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Mahmood Moshfeghian

Reference:
1. Campbell, J.M., Gas Conditioning and Processing, Volume 1: The Basic Principles, 9^th Edition, 2^nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
2. Moshfeghian, M., Maddox, R.N., and A.H. Johannes, “Application of Exponential Decay Distribution of C₆₊ Cut for Lean Natural Gas Phase Envelope,” J. of Chem. Engr. Japan, Vol 39, No 4, pp.375-382 (2006)
3. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
4. Soave, G., Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.
5. Brill, J. P., et al., “Analysis of Two-Phase Tests in Large-Diameter Flow Lines in Prudhoe Bay Field,” SPE Jour, p. 363-78, June 1981.
Figure 6. Variation of oil and gas velocity for two stock tank oil rate along the gathering line at 6900 kPag for 101.6 mm pipe diameter and GOR of 89 Sm³/STm³ (500 scf/STB)

Figure 7. Impact of GOR, Sm³/STm³ (scf/STB) on pressure drop along the gathering line at 6900 kPag for 101.6 mm pipe diameter
November 1, 2014
Impact of Gas-Oil Ratio (GOR) on Crude Oil Pressure Drop in Gathering Systems
The use of multiphase flow systems is common practice in the oil and gas industry. Multiphase flow is often encountered in the well tubing, flow lines and gathering systems. For transport of oil and gas (and water) to downstream processing facilities the preference is normally a single pipeline in which both phases are transported simultaneously for economic reasons. Even in gas pipelines where the gas enters the line as a single phase fluid, condensation of liquids can occur due to pressure and temperature changes along the line.

Modeling and simulation of a multiphase systems, even under steady-state conditions, is complex. There are a few tools designed specifically for modeling and analysis of complex multiphase systems such as PipePhase, PipeSim, OLGA, etc. [1].

In the June 2008 Tip of the Month (TOTM), we demonstrated how general-purpose process simulation programs can be used to simulate gas dominated two-phase pipelines. In the August 2008 TOTM, we discussed the value of the simple Flanigan correlation and how it can be used to model and analyze the behavior of a wet gas transmission pipeline. The results of the Flanigan correlation were compared with more rigorous calculation methods for multiphase pipelines.

In this TOTM, we will study the impact of gas-oil ratio (GOR) on pressure drop in crude oil gathering systems. Specifically, pressure drop along a gathering line for nominal pressures of 690, 3450, and 6900 kPag (100, 500, and 1000 psig) and nominal pipe size of 101.6 and 152.4 mm (4 and 6 inches) was calculated using multiphase rigorous method from commercial simulator. The calculated pressure drops are presented in graphical format as a function of the oil stock tank volume flow rate and GOR. Variation of thermo physical properties was considered.

Case Study

For the purpose of illustration, we considered a case study for transporting a crude oil of relative density of 0.852 (°API = 34.6) at stock tank condition combined with a gas with relative density of 0.751. The selected GORs were 0 (dead oil), 17.8, 356.5, and 891.3 Sm³ of gas/STm³ of oil (0, 100, 2000, and 5000 scf/STB). The compositions of oil and gas are presented in Table 1. The oil C₆₊ was characterized as 30 single carbon number (SCN) [2] ranging from SCN6 to SCN35 while the gas C₆₊ was characterized by 10 SCN ranging from SCN6 to SCN15. For details of the SCN components, see Table 3.2 on page 64 of reference [2]. The mole fraction of SCN components were determined by an exponential decay algorithm [3].

Table 1. Feed composition at stock condition

The following assumptions were made:
1. Steady state conditions
2. The line is 1.601 km (1 mile) long with nominal size of 101.6 and 152.4mm (4 and 6 inches), onshore buried line.
3. Segment lengths and elevation changes are presented in Table 2. This elevation profile is considered to be approximately equivalent to “rolling” terrain.
4. Pipeline inside surface roughness of 46 microns (0.046 mm, 0.0018 inch)
5. Line nominal pressure 690, 3450, and 6900 kPag (100, 500, and 1000 psig)
6. The feed enters the line at 15.6 ˚C and (60 ˚F)
7. The ground/ambient temperature, is 15.6 ˚C and (60 ˚F)
8. Water cut is 0 (no water in the feed).
9. Overall heat transfer coefficients of 2.839 W/m²-˚C (0.5 Btu/hr-ft²-˚F), for onshore buried line (minor effect as inlet temperature = ambient ground temperature).
10. Simulation software ProMax [4] and using the Soave-Redlich-Kwong (SRK) Equation of State [5] for vapor-liquid equilibrium and Beggs-Brill method for two-phase pressure drop calculation [6].
Table 2. Line segment length and elevation change

Results and Discussions:

The two phase (oil and gas) flow through the gathering line was simulated by ProMax with SRK EOS for vapor-liquid equilibria and Beggs-Brill for two phase pressure drop calculations. Figures 1A and 1B present the calculated pressure drop per unit length as a function of oil stock tank volume rate and GOR for nominal line diameter of 101.6 mm (4 inches) at nominal line pressure of 690 kPag (100 psig) in SI (international) and FPS (Engineering) system of units, respectively. Figures 1A and 1B indicate that as the GOR increases from 0 to 891 Sm³/STm³ (0 to 5000 scf/STB), the pressure drop increases considerably. Consequently, as the GOR increases, the line capacity decreases.

Figures 2A, 2B, 3A, and 3B present the results for the same line size but at nominal pressures of 3445 and 6900 kPag (500 and 1000 psig), respectively. Contrary to Figure 1, Figures 2 and 3 indicate that at these higher pressures as the GOR increases, the pressure drop decreases for low GOR value. However, for further increase of GOR the pressure drop increases considerably.

Similar calculations were repeated for another line with nominal pipe size of 152.4 mm (6 inches) and the simulation results are presented in Figures 4 through 6. Figures 4 through 6 also demonstrate the same impact of GOR on the pressure drop, at higher pressures and low GOR, the pressure drop decreases. However, the impact of low GOR at higher pressures is less compared to the smaller line diameter.

Figure 1A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 690 kPag for 101.6 mm pipe diameter

Figure 1B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 100 psig for 4 in pipe diameter

Figure 2A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 3445 kPag for 101.6 mm pipe diameter

Figure 2B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 500 psig for 4 in pipe diameter

Figure 3A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 6900 kPag for 101.6 mm pipe diameter

Figure 3B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 1000 psig for 4 in pipe diameter

Conclusions

The following conclusions can be made based on this case study:
1. The GOR has a large impact on the capacity of crude oil gathering lines. In general as GOR increases the pressure drop increases which lowers the line capacity.
2. At high pressures and low GOR, pressure drop is lower than the pressure drop for dead oil (solution gas is zero) because the viscosity of live oil is lower than viscosity of dead oil. This effect is bigger for the smaller line diameter.
To learn more about similar cases and how to minimize operational problems, we suggest attending our PF 45 (Onshore Gas Gathering Systems: Design and Operation), G4 (Gas Conditioning and Processing), PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Mahmood Moshfeghian

Reference:
1. Ellul, I. R., Saether, G. and Shippen, M. E., “The Modeling of Multiphase Systems under Steady-State and Transient Conditions – A Tutorial,” The Proceeding of Pipeline Simulation Interest Group, Paper PSIG 0403, Palm Spring, California, 2004.
2. Campbell, J.M., Gas Conditioning and Processing, Volume 1: The Basic Principles, 9^th Edition, 2^nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
3. Moshfeghian, M., Maddox, R.N., and A.H. Johannes, “Application of Exponential Decay Distribution of C₆₊ Cut for Lean Natural Gas Phase Envelope,” J. of Chem. Engr. Japan, Vol 39, No 4, pp.375-382 (2006)
4. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
5. Soave, G., Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.
Brill, J. P., et al., “Analysis of Two-Phase Tests in Large-Diameter Flow Lines in Prudhoe Bay Field,” SPE Jour, p. 363-78, June 1981.

Figure 4A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 690 kPag for 152.4 mm pipe diameter

Figure 4B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 100 psig for 6 in pipe diameter

Figure 5A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 3445 kPag for 152.4 mm pipe diameter

Figure 5B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 500 psig for 6 in pipe diameter

Figure 6A (SI). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 6900 kPag for 152.4 mm pipe diameter

Figure 6B (FPS). Variation of pressure drop per unit length with oil stock tank volume rate and GOR at 1000 psig for 6 in pipe diameter
October 1, 2014
Lean Sweet Natural Gas Water Content Correlation
In the October, November, December 2007 and February 2014 Tips of the Month (TOTM), we studied in detail the water phase behaviors of sweet and sour natural gases and acid gas systems. We also evaluated the accuracy of different methods for estimating the water content of sour natural gas and acid gas systems.

The water vapor content of natural gases in equilibrium with water is commonly estimated from Figure 6.1 of Campbell book [1] or Figure 20.4 of Gas Processors and Suppliers Association, including corrections for the molecular weight (relative density) of gas and salinity of water [2].

In this TOTM, we will present two new correlations for estimating the water content of lean and sweet natural gases. The performance of the proposed correlations will be compared with the rigorous simulation and shortcut method software and other correlations.

Low Pressure System

At low pressure conditions, less than 700 kPa (100 psia), the mole fraction of water in the gas phase can be estimated by dividing water vapor pressure, P^V, at the specified temperature, T, by the system pressure, P. The vapor pressure of pure water, from 0 to 360, (32 to 680) can be calculated by the following relation [3].

Where:

The critical temperature, T_C = 647.096 K and critical pressure, P_C = 22.064 MPa, T in K, and P^V in MPa, and

a₁ = −7.85951783,     a₂ = 1.84408259,     a₃ = −11.7866497,      a₄ = 22.6807411,

a₅ = −15.9618719,     a₆ = 1.80122502

Knowing one kmole of water = 18 kg (lbmole=18 lb_m) and one kmole of gases occupy 23.64 Sm³ at standard condition of 15 and 101.3 kPa (one lbmole of gases occupy 379.5 SCF at standard condition of 60 and 14.7 psia), the water content is calculated by

Moderate to High Pressure System

For pressures higher than 700 kPa (100 psia), we propose a correlation similar to equation 6-276 in Chapter 6 of Standard Petroleum Handbook [4] as follows:

Reference [4] presents the tabular values of A and B as a function of temperature. In this work, the temperature dependency of A and B in equation 3 is presented in the form of Gaussian Model.

Where:

A temperature range of -40 to 100 (-40 to 212) and pressure range of 6.8 to 680 atm have been considered. For the purpose of higher accuracy, the parameters in equations 4G are regressed and presented in Table 1 for 5 different temperature intervals, for SI (System International) and FPS (Foot-pound-Second) system of units. The temperature range in the 5^th row in each system of units is used more frequently. The proposed correlations are suitable for spreadsheet calculations.

  Table 1. The parameters for equation 4G (Gaussian Model).

Alternatively, the temperature dependency of parameters A and B in equation 3 can also be presented in the form of Polynomial Model.

Similarly, the correlation parameters of equations 4P are presented in Table 2.

The performance of the proposed correlation was evaluated against the water content of a lean sweet natural gas with a relative density (specific gravity) of 0.6 calculated by SRK equation of state of ProMax software [5]. The water saturator tool of ProMax was utilized. Figure 1A (SI) and Figure 1B (FPS) present the comparison between the proposed correlation (Eqs. 3 and 4G) results (solid curves) and the corresponding results from ProMax designated by geometric symbols. Figure 1 indicates there is relatively good agreement between the proposed correlation and ProMax. Large deviations are observed for temperatures below -20 (-4). Almost in all conditions, except at very high temperature the proposed correlation is more conservative (over predicting) than the ProMax.

Table 2. The parameters for equation 4P (Polynomial Model).

Evaluation of the Proposed Correlation

The performance of the proposed correlation was also evaluated against the water content of a lean sweet natural gas predicted by GCAP software [6]. The water content of GCAP is based on Figure 6.1 of Campbell book [1]. Figure 2A (SI) and Figure 2B (FPS) present the comparison between the proposed correlation (Eqs. 3 and 4G) results (solid curves) and the corresponding results from GCAP designated by geometric symbols. Figure 2 indicates there is better agreement between the proposed correlation and GCAP compared to ProMax. Large deviations are observed for temperatures below -20 (-4). Almost in all conditions, except at very high temperature, the proposed correlation is slightly more conservative (over predicting) than the GCAP.

Figures 1 and 2 indicate that in the temperature range of 0 to 70 (32 to 158 ) excellent agreement is observed between results of equation 3 and both software. Figure 2 also indicates that the GCAP results for 170 atm at very low temperature is inconsistent with other pressure results of GCAP.

When equation 4P and the parameters presented in Table 2 were used instead of equation 4G, similar quality of results was obtained.

Figure 1A (SI). Comparison of results between the proposed correlation (Eq 3) and ProMax

Figure 1B (FPS). Comparison of results between the proposed correlation (Eq 3) and ProMax

Figure 2A (SI). Comparison of results between the proposed correlation (Eq 3) and GCAP

Figure 2B (FPS). Comparison of results between the proposed correlation (Eq 3) and GCAP

Bukacek Correlation

Bukacek [7] suggested a relatively simple correlation for the water content of lean sweet gas as follows:

where T is in °F.

This correlation is reported to be accurate for temperatures between 15.5 and 238°C (60 and 460°F) and for pressure from 0.105 to 69.97 MPa (15 to 10,000 psia). The pair of equations in this correlation is simple in appearance. The added complexity that is missing is that it requires an accurate estimate of the vapor pressure of pure water. In this study, we have used equation 1 for water vapor pressure.

The performance of the proposed correlation was also evaluated against the Bukacek’s Correlation coupled with equation 1 for water vapor pressure. Figure 3A (SI) and Figure 3B (FPS) present the comparison between the proposed correlation (Eqs. 3 and 4G) results (solid curves) and the corresponding results from Bukacek’s correlation designated by geometric symbols. Figure 3 indicates there is an excellent agreement between the proposed correlation and Bukacek’s correlation.

Conclusions

The following conclusions can be made regarding the proposed correlation:
1. A relatively simple correlation for predicting the water content of lean sweet natural gas is presented that can be used for spreadsheet calculation.
2. Based on Figures 1 and 2, the agreement between the predicted water content by the proposed correlation (Eqs. 3 and 4G or Eqs. 3 and 4P) and those predicted by ProMax and GCAP software is relatively good. The agreement deteriorates at temperatures below -20 (-4).
3. In general, the estimated water content by the proposed correlation is conservative compared to ProMax and GCAP.
4. A better agreement between the proposed correlation and GCAP compared to ProMax is observed.
5. Excellent agreement is observed between the proposed and the Bukacek’s correlations.
6. The GCAP results for pressure of 170 atm at low temperatures are inconsistent with respect to GCAP results at other pressure.
7. The proposed correlation is easy and suitable for hand or spreadsheet calculations.
To learn more about similar cases and how to minimize operational problems, we suggest attending our G6 (Gas Treating and Sulfur Recovery ), G4 (Gas Conditioning and Processing ), PF81 (CO₂ Surface Facilities ), PF4 (Oil Production and Processing Facilities ), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Dr. Mahmood Moshfeghian

Reference:
1. Campbell, J.M., Gas Conditioning and Processing, Volume 1: The Basic Principles, 9^th Edition, 2^nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
2. GPSA Engineering Data Book, Section 20, Volume 2, 13^th Edition, Gas Processors and Suppliers Association, Tulsa, Oklahoma, 2012.
3. Wagner, W. and Pruss, A., J. Phys. Chem. Reference Data, 22, 783–787, 1993.
4. Standard Handbook of Petroleum, Natural Gas Engineering volume 2, Lyons, W. C., Editor, Gulf Professional Publishing, Houston, Texas, 1996
5. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
6. GCAP 9.1, Gas Conditioning and Processing, PetroSkills/Campbell, Norman, Oklahoma, 2014
7. Bukacek, R.F., “Equilibrium Moisture Content of Natural Gases” Research Bulletin IGT, Chicago, vol 8, 198-200, 1959.
Figure 3A (SI). Comparison of results between the proposed correlation (Eq 3) and Bukacek’s Correlation

Figure 3A (FPS). Comparison of results between the proposed correlation (Eq 3) and Bukacek’s Correlation
September 1, 2014
Gas Sweetening-Part 1: Comparison of Amines
Hydrogen sulfide and carbon dioxide are the principal objectionable acid gas constituents often present in natural gas, synthetic gas, and various refinery gas streams. These acid gas constituents must be removed for corrosion prevention in gas pipelines and process equipment and for health and safety reasons. Reference [1] provides current acceptable concentration levels for these acid gases in gas streams. Hydrogen sulfide removal is also often important for production of sulfur, which is used to create sulfuric acid and fertilizers. Carbon dioxide removal is also important for its capturing and sequestering, as well as for enhanced oil recovery.

In natural gas treating, there are several processes available for removing the acid gases. Aqueous solutions of alkanolamines are the most widely used [1]. The alkanolamines process is characterized as “mass transfer enhanced by chemical reactions” in which acid gases react directly or react through an acid-base buffer mechanism with an alkanolamines to form nonvolatile ionic species. For further detail of sour gas treating refer to references [1-4].

Several alkanolamines have been used for acid gas removal from natural gas streams. In this study only a primary monoethanolamine (MEA), a secondary diethanolamine (DEA) and a tertiary methydeithanolamine (MDEA) are considered. MEA has the highest reactivity and MDEA has the highest selectivity.

In this TOTM, we will study and compare the performance of these three amines by simulation of a simplified process flow diagram for removal of H₂S and CO₂ from a sour gas stream. The H₂S and CO₂ concentration in the sweet gas, amine solution circulation rate, reboiler duty, amine losses, pump power, and lean-rich heat exchanger (HEX) duty are calculated and plotted for a wide range of steam rates needed to regenerate the rich solution. In addition, the optimized steam rates and corresponding design variables are determined and reported.

Case Study:

For the purpose of illustration, we considered sweetening of 1.416×10⁶ std m³/d (50 MMSCFD) of a sour and wet natural gas with the composition, pressure, and temperature presented in Table 1. ProMax [5] simulation software with “Amine Sweetening – PR” property package was used to perform all of the calculations.

Table 1. Feed composition, volumetric flow rate and conditions

The following specifications/assumptions were made:

Contactor Column
1. Feed sour gas is saturated with water
2. Number of theoretical stages = 8
3. Pressure drop = 35 kPa (5 psi)
4. Lean amine solution temperature = Sour gas feed temperature + 5.5 (10)
Regenerator Column
1. Number of theoretical stages = 11 (excluding condenser and reboiler)
2. Rich solution feed temperature = 98.9 (210)
3. Rich solution feed pressure = 245 kPa (35 psi)
4. Condenser temperature = 48.9 (120)
5. Pressure drop = 35 kPa (5 psi)
6. Bottom pressure = 138 kPa (20 psig)
7. Reboiler duty = Specified “Steam Ratio” times circulation rate (Refer to Table 2)
Heat Exchangers
1. Lean amine cooler pressure drop = 21 kPa (3 psi)
2. Rich side pressure = 41 kPa (6 psi)
3. Lean side pressure = 35 kPa (5 psi)
Pump
1. Discharge Pressure = Sour gas feed pressure + 35 kPa (5 psi)
2. Efficiency = 65 %
Lean Amine Circulation Rate and Concentration
1. Varied to meet the target acid gas loading in the rich solution shown in Table 2
Rich Solution Expansion Valve
1. Pressure drop in first expansion valve (vlve 100) = 6310 kPa (915 psi)
2. Pressure drop in the second expansion valve (vlve 101) = 303 kPa (44 psi)
A simplified process flow diagram for the case studied is presented in Figure 1 [1].

Table 2. Specified amine concentration, target rich solution acid gas loading, and steam ratios [6

”]Results and Discussions:

Based on the description and specifications presented in the previous section, the process flow diagram in Figure 1 was simulated by ProMax [5]. The simulation was performed for steam ratios presented in Table 2. For each steam ratio and each amine, the H₂S and CO₂ concentration in the sweet gas, lean amine circulation rate, reboiler duty, and the amine make up to compensate the losses due to vaporization from top of contactor and regenerator columns, and flashed gas in the separator are calculated. The variation of these properties as a function of steam rate is presented in Figures 2 through 8.

Figure 2. Sweet gas H2S content vs steam rate

Figure 3. Sweet gas CO2 content vs steam rate

Figures 2 and 3 present the variation of H₂S and CO₂ concentration in the sweet gas stream as a function of steam rate for MEA, DEA, and MDEA. Figure 2 also indicates that the minimum steam rate to achieve common pipeline specification of H₂S concentration of 4 PPM. It should be noted that for the same H₂S concentration in the sweet gas, MDEA requires the lowest steam rate. Figure 3 indicates that both MEA and DEA do much better at removal of CO₂ than MDEA. Because MDEA requires the lowest steam rate, it is a preferred amine for selective removal of H₂S.

The required amine circulation rate as a function of steam rate is presented in Figure 4 for MEA, DEA, and MDEA. Figure 4 indicates that MDEA requires the smallest circulation rate for regeneration. In addition, the MDEA circulation rate is much lower than that of the other two amines. This is because MDEA has a much higher concentration (smaller amount of water) and can have higher maximum allowable acid gas loading in the rich solution (Refer to Table 2) compare to MEA and DEA.

The required reboiler duty as a function of steam rate is presented in Figure 5 for MEA, DEA, and MDEA. Figure 5 indicates that MDEA requires the smallest heat duty due to its very low circulation rate.

Figure 6 presents the amine make up as a function of steam rate for the three amines. This figure indicates that MEA has the highest and DEA the lowest vaporization loss. MDEA loss is between MEA and DEA because the normal boiling point of MDEA is between that of MEA and DEA (refer to Table 2). It should be noted that this loss does not include entrainment (mechanical) from the top of contactor. Amine vaporization loss from top of the regenerator column was practically zero for all three amines. The mechanical (entrainment) loss is normally much higher than vaporization loss.

Figures 7 and 8 present the required pump power and lean-rich amine heat exchanger duty as a function of steam rate for the three amines, respectively. These two figures also indicate the MDEA requires the lowest pump power and heat exchanger duty due to its lowest circulation rate.

Figure 4. Circulation rate vs steam rate

Figure 5. Reboiler duty vs steam rate

Figure 6. Total amine vaporization loss vs steam rate

Figure 7. Pump power vs steam rate

Figure 8. Lean-Rich HEX duty vs steam rate

Optimized Condition

For each amine the optimized/minimum steam rate to meet sweet gas H₂S content of 4 PPM was determined and the corresponding parameters are calculated and reported in Tables 3 and 4.

Table 3. Optimized parameters for three amines

Table 4 indicates that the optimized reboiler duty for DEA and MDEA are within the approximate guideline provided by the GPSA data book [3]; however, MEA reboiler duty is below the approximate guideline. According to the GPSA data book, reboiler duty varies with regenerator overhead reflux ratios, rich solution feed temperature to regenerator and reboiler temperature. In this case the same values of the above mentioned variables were used for the three amines. For a detailed comparison, for each amine the optimized variables should be selected.

Table 4. Comparison of reboiler duty with GPSA data book [3

Conclusions:

Based on the results obtained for the case study considered in this TOTM, the following conclusions can be made:
1. MDEA is selective in removal of H₂S and allows some of the CO₂ to slip through (Figures 2 and 3).
2. For a specified H₂S content in the sweet gas, regeneration of MDEA requires:
  1. the lowest steam rate (reboiler duty).
  2. the lowest pump power.
  3. the lowest lean-rich HEX duty.
3. MDEA vaporization loss is between MEA and DEA.
4. Amine vaporization loss from the top of the regenerator column is practically zero.
5. The entrainment (mechanical) losses are much greater than the vaporization losses
To learn more about similar cases and how to minimize operational problems, we suggest attending our G6 (Gas Treating and Sulfur Recovery), G4 (Gas Conditioning and Processing), PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Dr. Mahmood Moshfeghian

Reference:
1. Maddox, R.N., and Morgan, D.J., Gas Conditioning and Processing, Volume 4: Gas treating and sulfur Recovery, Campbell Petroleum Series, Norman, Oklahoma, 1998.
2. Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9^th Edition, 1^st Printing, Editors Hubbard, R. and Snow –McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
3. GPSA Engineering Data Book, Section 21, Volume 2, 13^th Edition, Gas Processors and Suppliers Association, Tulsa, Oklahoma, 2012.
4. Moshfeghian, M., Bell, K.J., Maddox, “Reaction Equilibria for Acid Gas Systems, Proceedings of Lawrence Reid Gas Conditioning Conference, Norman, Oklahoma, 1977
5. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
6. Sour Gas Processing Training Manual, Bryan Research and Engineering, Inc., Bryan, Texas, 2014.
July 2, 2014
Refrigeration with Heat Exchanger Economizer vs Simple Refrigeration System
The details of a simple single-stage refrigeration system, a two-stage refrigeration system employing one flash tank economizer, and with heat exchanger economizer system are given in Chapter 15 of Gas Conditioning and Processing, Volume 2 [1].

In the January 2008 Tip of the Month (TOTM) [2], we compared the performance of a simple refrigeration system with another employing a flash economizer. Specifically, we evaluated compressor power saving, the effects of compressor suction–line pressure drop and the interstage pressure drop on compressor power requirement and condenser duty.

A second type of economizer configuration is the heat exchanger economizer shown in Figure 1, which is the same as Figure 15.9 of reference [1]. Cold, low-pressure chiller vapor is used to subcool the saturated liquid refrigerant. This decreases the refrigerant circulation rate, and may reduce compressor power. In this TOTM we will evaluate quantitatively the performance of a case study comparing a simple refrigeration system with another one containing heat exchanger economizer.

The process flow diagrams for the simple and with heat exchanger (HEX) economizer refrigeration systems are shown in Figure 2. Note that provisions have been made to consider pressure drop in different segments of the loops.

”]Let’s consider removing 1.0×10⁷kJ/h which is equal to 2778 kW (9.479 MMbtu/hr) from a process gas at -35°C (-31°F) and rejecting it to the environment by the condenser at a condensing temperature of 35°C (95°F). Assuming 5 kPa (0.7 psi) pressure drop in the chiller and 5 kPa in the suction line pressure drop, the compressor suction pressure is 132.4 kPa (19.1 psi). The condenser pressure drop plus the pressure drop in the line from the compressor discharge to the condenser was assumed to be 50 kPa (7.3 psi); therefore, compressor discharge pressure is 1270 kPaa (184.2 psia). The compressor discharge temperature is 66°C (150.8°F). At these conditions, the condenser duty is 4434 kW (15.13 MMbtu/hr). Pure propane is used as the working fluid. In this study, all of the simulations were performed by UNISIM software [3].

Figure 2. Process flow diagrams for a simple refrigeration system and with a heat exchanger economizer

In order to study the effect of HEX economizer, we considered the following scenarios:
1. The condensed liquid at temperature of 35°C (95°F) was cooled starting from 33 to 24 °C (91.4 to 75.2°F) with a step change of -1°C (-1.8°F).
2. Step 1 was repeated three times for pressure drops of 20, 25, and 30 kPa ( 2.9, 3.63, or 4.4 psi) on both sides of HEX economizer.
3. For the above four cases the following variables were calculated:
4. The required compressor power
5. Compressor suction temperature
6. Compressor discharge temperature
7. Refrigerant (propane) circulation rate
8. Condenser heat duty
9. Chiller inlet temperature
Figures 3, 4, and 5 present the required compressor power, condenser duty, and HEX duty as a function of liquid propane subcooled temperature (at the outlet of HEX economizer), respectively. Figures 3 and 4 indicate that as the propane subcooled temperature decreases the compressor power and condenser duty decrease, too. However as the pressure drop in the cold vapor (low pressure) side increases, the compressor power and condenser duty increase. The pressure drop significantly increases the compressor power. Figure 5 indicates as the propane subcooled temperature decreases, the HEX duty increases independent of HEX pressure drop.

Figure 3. Compressor power as a function of refrigerant subcooled temperature and HEX economizer pressure drop

Figure 4. Condenser duty as a function of refrigerant subcooled temperature and HEX economizer pressure drop

Figure 5. HEX duty as a function of refrigerant subcooled temperature and HEX economizer pressure drop

The refrigerant mass circulation rate, compressor suction temperature, and compressor discharge temperature as a function of propane subcooled temperature and HEX pressure drop are presented in Figures 6, 7, and 8, respectively. Figure 6 indicates as the propane subcooled temperature decreases, the mass circulation rate decreases independent of HEX pressure drop. Figures 7 and 8 indicate that compressor suction and discharge temperatures increase with decrease in propane subcooled temperature. However, the effect of HEX pressure drop on discharge temperature is more pronounced.

Contrary to a refrigeration system with a flash thank economizer in which the compressor power is reduced [2], by employing HEX economizer the power requirement increased. With regard to the compressor power, two factors offset the reduced circulation rate. The first is HEX pressure drop. The pressure drop on the low pressure side significantly increased compressor power, since the suction pressure was near atmospheric. Secondly, the refrigerant vapor entering the compressor is now super-heated. Although this reduces the likelihood of liquid carryover into the compressor, it resulted in higher power consumption due to the higher suction temperature.

To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), and PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Dr. Mahmood Moshfeghian

Figure 6. Refrigerant mass circulation rate as a function of refrigerant subcooled temperature and HEX economizer pressure drop

Figure 7. Compressor suction temperature as a function of refrigerant subcooled temperature and HEX economizer pressure drop

Figure 8. Compressor discharge temperature as a function of refrigerant subcooled temperature and HEX economizer pressure drop

Reference:
1. Campbell, J.M., “Gas conditioning and Processing, Vol 2: The Equipment Modules”, 9^th Edition, Edited by R.A. Hubbard K.S. McGregor, John M. Campbell & Company, Norman, USA, 2014.
2. Moshfeghian, M., “Refrigeration with Flash Economizer vs Simple Refrigeration System,” http://www.jmcampbell.com/tip-of-the-month/2008/01/refrigeration-with-flash-economizer-vs-simple-refrigeration-system/ , 2008
3. UniSim Design, Version 410 Build 17061, Honeywell International, Inc., Calgary, Canada, 2013.
May 1, 2014
Transportation of Ethane by Pipeline in the Dense Phase
In the January, February, and March 2012 tips of the month (TOTM) we discussed the transportation of carbon dioxide (CO₂) in the dense phase region. We illustrated how thermophysical properties changed in the dense phase and studied their impacts on pressure drop calculations. We showed that the effect of the numerical range of values for the overall heat transfer coefficient on the pipeline temperature is significant.

In this TOTM, we will study the transportation of ethane by pipeline in the dense phase region. For a case study, a mixture of ethane containing a small fraction of methane and propane was considered. The pressure and temperature profiles along the pipelines are calculated and plotted on the feed phase envelope. In addition, the pump power requirement, pressure and temperature profiles for a single pipeline with a lead pump station are compared with the option of dividing the same line into three equal segments having one lead pump station and two intermediate pump stations.

Calculation Procedure:

For a pure compound above critical pressure and critical temperature, the system is often referred to as a “dense fluid” or “super critical fluid” to distinguish it from normal vapor and liquid (see Figure 1 for carbon dioxide in December 2009 TOTM [1]).

The same step-by-step calculation procedure described in the February 2012 TOTM [2] was used to determine the pressure and temperature profiles in a pipeline.

In the following section we will illustrate the pressure drop calculations for transporting ethane mixture in the dense phase. For details of pressure drop equations in the gas and liquid phases refer to the January 2012 TOTM [3].

Case Study:

For the purpose of illustration, we considered a case study for transporting 31393 kg/h (69209.0 lb_m/hr) equivalent to 275000 tonne/y of the cited ethane mixture with the composition presented in Table 1. The mixture is available at the pressure and temperature presented in Table 1.

The following assumptions were made:
1. Horizontal pipeline, no elevation change
2. Steady state conditions
3. The pipeline is 1380 km (858 mile) long with an inside diameter of 208.6 mm (8.212 in), onshore buried line.
4. Pipeline inside surface roughness of 46 microns (0.046 mm, 0.0018 inch) which is equivalent to inside surface relative roughness (roughness factor), ε/D, of 0.00022.
5. Delivery Pressure is 5.3 MPa (769 psia)
6. The ground/ambient temperature, is 18.3 ˚C, (65 ˚F)
7. Overall heat transfer coefficient of 2.839 W/m²-˚C (0.5 Btu/hr-ft²-˚F), for onshore buried line.
8. Pump efficiency is 50%, this is the worst case, the actual pump efficiency is in the range of 50-85%).
9. Simulation software: ProMax [4] and using Soave-Redlich-Kwong (SRK) [5] Equation of State.
The block diagrams for two options studied are presented in Figure 1. In option A, only one pump station and a single long segment were considered. In option B, for the same pipeline and feed conditions, one lead and two intermediate pump stations with three equal pipeline segment were considered. Each segment length is 460 km (286 mile), which is 1/3 of total length and all have the same inside diameter of 208.6 mm (8.212 in). The delivery pressures for both options are the same.

Results and Discussions:

Figure 2 presents the phase envelope for the ethane mixture with the composition presented in Table 1. According to Figure 2, the feed at the pump suction conditions of 0°C (32°F) and 3000 kPa (435 psia) presented in Table 1 is in the liquid phase. In order to deliver the ethane mixture at pressure of 5.3 MPa (769 psia), for option A this liquid is pumped to a pressure of 13.6 MPa (1972 psia) before entering the pipeline. Due to pumping, the feed temperature rises from 13.6 MPa (1972 psia) before entering the pipeline. A step-by-step pump calculation with increments of 2 MPa (290 psia) for the discharge pressure reveals that the temperature rise is linear with pressure. This significant temperature rise is due to compressibility of ethane mixture.

Figure 2. Phase envelope for ethane mixture.

The pumping and pipeline pressure-temperature paths for option A are plotted on the phase envelope and presented in Figure 3. As shown in this figure, the ethane mixture at the pump discharge (pipeline inlet) is in the supercritical region (dense phase). Figure 3 indicates that as the feed enters the pipeline, its temperature drops rapidly and remains constant and very close to the ambient temperature.

Figure 3. Phase envelope, pumping path and dense phase pipeline pressure-temperature profile.

Figure 4 presents the calculated pipeline pressure profiles for options A and B. For option A, the inlet pressure is 13.6 MPa (1972 psia) to assure ethane mixture delivery at 5.3 MPa (769 psia). Similarly, in option B at each pump station the pressure is increased to 8.2 MPa (1189 psia).

Figure 5 presents the calculated pipeline temperature profiles for options A and B. The constant ambient temperature of 18.3˚C and (65˚F) is also plotted. In option B, the discharge temperature for lead pump station is 11.2˚C, (52.1˚F) which is below the ambient temperature. For both options, the pipeline temperature rapidly approaches the ambient temperature within the first 50 km (31 mile).

Figures 6 and 7 present the density and velocity profiles along the pipeline, respectively. For a crude oil cross-country pipeline, the velocity is in the range 1.5 to 2.5 m/s (5 to 8 ft/sec). The abrupt change of density, and consequently velocity, along the first 60 km (37.3 mile) is due to the ethane mixture temperature drop, which approaches the ground temperature.

Figure 4. Pipeline pressure profiles for options A and B.

Figure 5. Pipeline temperature profiles for options A and B.

Figure 6. Pipeline fluid density profile for option A

Figure 7. Pipeline fluid velocity profile for option A

The total pump power requirements with pump efficiency of 50% for options A and B are 457 and 504 kW (381.6 and 420.8 hp), respectively. This is for screening estimate only. Normally, work is done in collaboration with the pump manufacturers for better efficiencies based on CAPEX and OPEX.

Conclusions:

Based on the results obtained for the case study considered in this TOTM, the following conclusions can be made:
1. During the pumping of ethane mixture, the temperature rises linearly with pressure (Figure 2).
2. The feed temperature approaches the ground temperature rapidly (Figure 5). This may not be the case for lower overall heat transfer coefficient.
3. A single pipeline with only one lead pump station (option A) requires smaller pump power compared to the option of one lead pump and two intermediate pump stations (option B). Due to higher pressure in option A, wall thickness will be higher.
4. A complete cost analysis of pumping requirement vs pipeline cost should be made to determine the optimum pipeline diameter, wall thickness and power requirement.
5. The point not considered but worth mentioning is that ethane is very difficult to seal. We would work with pump and seal manufacturers for selecting the correct dry gas seal. This selection could determine the overall system reliability.
To learn more about similar cases and how to minimize operational problems, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), and PL4 (Fundamentals of Onshore and Offshore Pipeline Systems) courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Dr. Mahmood Moshfeghian

Reference:
1. Bothamley, M.E. and Moshfeghian, M., “Variation of properties in the dese phase region; Part 1 – Pure compounds,” TOTM, http://www.jmcampbell.com/tip-of-the-month/2009/12/variation-of-properties-in-the-dense-phase-region-part-1-pure-compounds/, Dec 2009.
2. Moshfeghian, M., ”Transportation of CO₂ in the Dense Phase,” TOTM, http://www.jmcampbell.com/tip-of-the-month/2012/02/ , Feb 2012
3. Moshfeghian, M., ”Transportation of CO₂ in the Dense Phase,” TOTM, http://www.jmcampbell.com/tip-of-the-month/2012/01/, Jan 2012
4. ProMax 3.2, Bryan Research and Engineering, Inc, Bryan, Texas, 2014.
5. Soave, G., Chem. Eng. Sci. 27, 1197-1203, 1972.
March 1, 2014
Acid Gas-Water Content
In the past Tips of the Month (TOTM), we discussed the phase behavior of sweet natural gas- water, sour natural gas-water, and acid gas–water systems. They were posted in October 2007 TOTM [1], November 2007 TOTM [2], and December 2007 TOTM [3], respectively. In this TOTM, we will revisit the acid gas-water phase behavior system. Specifically, different methods of predicting water content of acid gas systems are evaluated based on experimental data from the literature. Water content diagrams compatible with the experimental data for pure CO₂, Pure H₂S, pure CH₄ and their mixtures are generated and presented. These charts can be used for facility type calculations and trouble shooting.

Table 1 presents the compositions of several acid gas mixtures evaluated in this study, along with their saturated water contents (in mole percent) from experimental data [4] and from predictions by five methods. The Maddox et al. [5] results were generated using GCAP [6] software and the Erbar et al. [7] results were generated by EzThermo [8] software. The Wichert & Wichert [9] and Yarrison et al. [10] results are from GPA RR-210 [11]. The last column presents the results predicted by SRK in ProMax [12].

Table 1 indicates that as long as the total acid gas concentrations is less than 60 mole percent, all five methods produce results within the accuracy of experimental data. However, for higher concentrations of acid gases, the Yarrison et al. [10] and ProMax [12] methods provide more accurate results.

It should be noted that the results of ProMax in Table 1 are based on the water saturator tool available in ProMax. If a conventional 3-phase separator calculation of ProMax is used, the error percent for the third row from the bottom of table reduces from -44.9 % to -16.6 %. The results for the other cases remained practically the same for these two calculations options.

Table 2 and Figure 1 present the error analysis for prediction of 74 additional gas mixtures containing acid gases. The detail of data points and sources of experimental points are in GPA RR-210 [11]. The five methods under study are Maddox et al. [5], Robinson et al. [13], Wichert & Wichert [9], Yarrison et al. [10], and ProMax [12]. With the exception of the ProMax results, the predicted water contents used for error analysis of the other four methods were extracted from GPA RR-210 [11].

Figure 1 presents the error analysis graphically for the same 74 gas mixtures presented in Table 2. Based on the error analysis of Tables 1, 2, and Figure 1, the ProMax method was chosen to generate water content diagrams for pure CO₂, pure H₂S, and their mixtures. These diagrams are presented in the proceeding sections.

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. Figures 2 (SI) and 2 (FPS) present the water content of pure H₂S predicted by ProMax [12] as a function of pressure and temperature, in the international system (SI) and engineering system of units (FPS, foot, pound and second). The behavior shown on this plot is quite complicated and explained thoroughly by Carroll [14]:

“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.”

Within the transition region, the acid gas exists as both liquid and vapor. Water saturation of the vapor phase is represented by the lower value, whereas the water content of the liquid phase is the higher value.

Figures 3 (SI) and 3 (FPS) present the water content of pure CO₂ and pure CH₄ predicted by ProMax [12] as a function of pressure and temperature, in the international system (SI) and engineering system of units (FPS). When Figures 2 and 3 were superimposed on Figures 20-5 and 20-6, respectively, of the GPSA data book [15] a very good match was obtained. The two figures in the GPSA data book are based on experimental data and the Yarrison et al. model.

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 32.2°C (90°F). As shown in Figure 3 (as well as in Figure 2 reported by Maddox and Lilly [16]), the water content of CO₂ exhibits a minimum.

Figure 4 presents the phase behavior of pure CO₂, Pure H₂S and three mixtures of them containing 2 mole percent CH₄. Their corresponding water content charts are presented in Figure 5.

Summary:

There are several methods available that can be used to predict the water content of acid gases. Most of these methods are based on equations 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. Although not addressed in this study, hydrates can also form and these can significantly complicate phase behavior.

Different methods of predicting water content of acid gas systems are evaluated based on the literature experimental data. In addition, the water content diagrams compatible with the experimental data for pure CO₂, H₂S, CH₄ and their mixture are generated and presented. These charts can be used for facility type calculations and trouble shooting.

To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), and PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

References:
1. Moshfeghian, M. “Water-Sweet Natural Gas Phase behavior,” http://www.jmcampbell.com/tip-of-the-month/2007/10/water-sweet-natural-gas-phase-behavior/, October 2007.
2. Moshfeghian, M., ”Water-Sour Natural Gas Phase Behavior,” http://www.jmcampbell.com/tip-of-the-month/2007/11/water-sour-natural-gas-phase-behavior/, November 2007.
3. Wright, W. and M. Moshfeghian, “Acid Gas-Water Phase Behavior,” http://www.jmcampbell.com/tip-of-the-month/2007/12/acid-gas-water-phase-behavior/, December, 2007.
4. 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.
5. 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.
6. GCAP 8.3, John M. Campbell & Co., Norman, Oklahoma, December 2010.
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. EzThermo, Chemical Engineering Consultants, Inc, Stillwater, Oklahoma, 2010.
9. Wichert, G. C. and E. Wichert, “New Charts Provide Accurate Estimation for Water Content of Sour Natural Gas”, O&G J, pp 64-66, Oct. 27, 2003..
10. Yarrison M., Song, K. Y., Cox, K,, Chronister D. and Chapman, W., “Water Content of High Pressure, High Temperature Methane, Ethane and Methane+CO2, Ethane + CO2,” RR-200, GPA, Tulsa, OK, March, 2008.
11. Song, K. Y., Vo, T., Yarrison M. and Chapman, W., “Acid Gas Water Content, An Update Of Engineering Data Book I,” RR-210, GPA, Tulsa, OK, June, 2012.
12. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, U.S.A., 2013.
13. Robinson, J. N., et al., Trans. AIME, Vol. 263, p. 281, 1977
14. 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.
15. GPSA Data Book, Vol. 2, 13^th Ed., Gas Processors and Suppliers Association, Tulsa, Oklahoma, 2013.
16. 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.
February 1, 2014
Estimating TEG Vaporization Losses in TEG Dehydration Unit
TEG Vaporization Losses

In this Tip of The Month (TOTM), the effect of striping gas rate and triethylene glycol (TEG) circulation ratio on the TEG vaporization loss from the regenerator top and contactor top is investigated. Specifically, this study focuses on the variation of TEG vaporization losses with reboiler pressure, TEG circulation ratio and stripping gas rate. By performing a rigorous computer simulation of TEG regeneration at reboiler pressures of 110.3 kPaa (16 psia) and 524.1 kPaa (76 psia), several charts for quick estimation of TEG vaporization losses from regenerator top and contactor, which are needed for facilities type calculations are developed. In addition, the effect of contactor temperature on the TEG vaporization losses for a case study is shown.

Computer Simulation Results:

In order to study the impact of the contactor temperature, stripping gas rate and TEG circulation rate on the TEG vaporization losses, the TEG dehydration process was simulated using ProMax [1] software with its Soave-Redlich-Kwong (SRK) [2] equation of state (EOS). The process flow diagram used for these simulations is the same as in the November 2013 TOTM [3] and is shown here in Figure 1.

The water-saturated gas with a water content of 915 mg/std m³ (57 lbm/MMSCF) enters the bottom of the contactor column at 37.8°C (100°F) and 6897 kPaa (1000 psia) at a rate of 2.835×10⁶ std m³/d (100 MMSCFD). The contactor column has three theoretical trays. The lean TEG solution enters at the top of the contactor column and flows down in the column. As shown in Figure 1, the water content of the dried gas is 10 mg/std m³ (0.63 lbm/MMSCF). The rich TEG solution contains 96.1 mass percent TEG entering the still column at 100°C (212°F) and 515 kPaa (74.7 psia). The reboiler temperature was set at 204.4°C (400°F) and boil-up ratio of 0.1 (molar bases). Two theoretical trays in the regenerator (still) column (N_R = 2) and two theoretical trays (N_S = 2) in the striping gas section were utilized. The striping gas enters the bottom of the stripping gas section at 204°C (399°F) and 524 kPaa (76 psia). Methane was used for the stripping gas at a rate of 56.3 std m³/h (1893 scf/hr). The regenerated lean solution contains 99.86 mass percent TEG and the ratio of stripping gas to lean TEG liquid volume rates is 20 std m³ of gas/std m³ of lean TEG solution (2.67 scf/sgal) or a mass ratio of 28.3. The regenerator (still) top temperature is 91.4°C (196.5°F). If the same stripping gas was sparged directly into the reboiler (N_S= 0, no stripping gas section), with everything else remaining the same, the regenerated solution contains 99.2 mass percent TEG and the regenerator column top temperature remains practically the same and is 91.1°C (196°F). For the above case the number of theoretical trays in the still column is increased from 2 to 3 (N_R = 3); the lean TEG concentration increased slightly from 99.6 to 99.8 mass percent but the regenerator column top temperature remained the same.

Using a similar set up as is shown in Figure 1, several simulations were performed for a range of stripping gas rates, for N_R=2, N_S=0 and for two reboiler pressures of 110.3 and 524 kPaa (16 and 76 psia) and temperature of 204.4°C (400°F). The results of these simulation runs are presented in Figures 2 to 5.

Figure 1. Sample results using ProMax [1] for TEG dehydration with reboiler P=110.3 kPaa (16 psia) with N_R=2 and N_S=2

The regenerator top temperatures are exactly the same as those in Figures 2 and 3 presented in the November 2013 TOTM [3].

Figure 2 presents the variation of the TEG vaporization losses from still/regeneartor column top with circulation ratio (mass basis) and stripping gas rate at top pressure of 101.3 kPaa (14.7 psia) and reboiler pressure of 110.3 kPaa (16 psia) operating at 204.4°C (400°F).

As was discussed in the November [3] and September [4] 2013 TOTMs, regeneration of TEG at higher reboiler pressure has several advantages such as preventing the emission of harmful contaminants like benzene, toluene, ethylbenzene, xylenes (BTEX), and hydrogen sulfide to the environment. Therefore, similar diagrams as shown in Figure 2 were generated for top pressure of 515.2 kPaa (74.7 psia) and reboiler pressure of 524.1 kPaa (76 psia) at 204.4°C (400°F). Figure 3 presents the variation of TEG vaporization loss from regenerator top for such a high reboiler pressure.

Fig 2. Variation of TEG vaporization loss from regenerator top with circulation mass ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F) and N_S=0

Fig 3. Variation of TEG loss from regenerator top with circulation msss ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F) and N_S=0

Figures 2 and 3 can be used for a quick estimate of the TEG vaporization loss from regenerator top for a given stripping gas rate and TEG circulation ratio either at low or high reboiler pressure. The two reboiler pressures selected in this study are typical operating pressures. For generation of data for Figures 2 and 3, the stripping gas was sparged directly into the reboiler; therefore, the number of theoretical trays for stripping gas section is zero (N_S=0). The corresponding figures in terms of TEG circulation volume ratio are presented in the Appendix (Figures 2A and 3A). Figures 2 and 3 indicate that as the stripping gas ratio increases the TEG vaporization losses decreases. These two figures also indicate that as the TEG mass circulation ratio increases, the TEG vaporization losses increases initially followed by a decreasing trend.

Generally, either 0, 1, or 2 ideal trays in the stripping gas section is used. In order to investigate the effect of the number of theoretical trays in the stripping gas section (N_S) on the TEG vaporization loss, simulations were performed for the cases of N_S=0 and N_S=2 for a constant stripping gas rate. Figure 4 presents the results of these simulations for low reboiler presssures of 110.3 kPaa (16 psia). The reboiler temperature for all cases was set at 204.4°C (400°F).

Figure 4 clearly indicates that the TEG vaporization loss from the regenerator top is independent of the number of ideal trays in the gas stripping section at low TEG mass circulation rates, however, it increases slightly with the increase in the number of ideal trays in the stripping gas section at higher TEG mass circulation ratio.

The effect of feed gas temperature to contactor and mass circulation ratio on TEG vaporization loss from the regenerator top is demonstrated in Figure 5. The TEG vaporization losses for three feed gas temperatures to contactor for stripping gas rate of 10 Std m³/m³ TEG (1.34 SCF/gal TEG) are plotted as a function of TEG mass circulation ratio. This figure indicates that as the contactor temperature increases, the TEG vaporization loss from the regenerator top increases. This can be explained because as the feed gas temperature increases, the feed gas water content (mass per unit volume at standard conditions) increases drastically which results in more vaporization of water from regenerator top. Consequently, more TEG (along with water) per unit volume of the gas at standard conditions is vaporized.

As expected Figure 6 indicates that the TEG vaporization loss from the contactor top is practically independent of the stripping gas rate. In addition this figure shows that as the TEG mass circulation ratio increases beyond 15 mass of TEG/mass of water removed, the TEG losses remain constant. As shown in Figure 7, Similar diagrams for higher reboiler pressure revealed almost the same observations.

Figure 8 also shows that as the TEG mass circulation ratio increases beyond 15 mass of TEG/mass of water removed, the TEG losses remain constant. However, as the feed gas temperature to the contactor increases, the TEG losses from the contactor top increase drastically.

Fig 4. Effect of number of ideal trays (N_S) in the gas stripping section on the TEG vaporization loss from regenerator top at P=101.3 kPaa (14.7 psia), reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F) and stripping gas rate of 10 Std m³/m³ TEG (1.34 SCF/gal TEG)

Fig 5. Effect of contactor temperature and mass circulation ratio on TEG vaporization loss from regenerator top at P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F), N_S=0 and stripping gas rate of 10 Std m³/m³ TEG (1.34 SCF/gal TEG)

Fig 6. Variation of TEG vaporization loss from contactor top with circulation mass ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F) and N_S=0

Fig 7. Variation of TEG vaporization loss from contactor top with circulation mass ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F) and N_S=0

Fig 8. Effect of contactor temperature and mass circulation ratio on TEG vaporization loss from contactor top at P=6897 kPaa (1000 psia), reboiler P=110 kPaa (16 psia) at 204.4°C (400°F) , N_S=0 and stripping gas rate of 10 Std m³/m³ TEG (1.34 SCF/gal TEG)

Conclusions:

In this TOTM, the effect of circulation ratio, stripping gas rate, theoretical number of trays, and the feed gas temperature to the contactor column on the TEG vaporization losses from contactor top and regenerator top for regeneration of TEG concentration at low and high reboiler pressure operating at 204.4°C (400°F) was studied. Charts for a quick estimation of the TEG vaporization losses from still/regenerator column top and contactor top at a specified stripping gas rate and circulation ratio to achieve a desired level of lean TEG concentration were prepared and presented in Figures 2, 3, 6, and 7. These charts are based on the rigorous calculations performed by computer simulations and can be used for facilities type calculations for evaluation and trouble shooting of an operating TEG dehydration unit. In addition, the following observations were made for the cases studied in this TOTM:
1. The TEG vaporization loss from the contactor top is almost 10 times higher than still/regenerator column top (see Figures 2, 3, 6 and 7).
2. As the feed gas temperature to the contactor column increases, the TEG vaporization loss from top of both columns increases (Figures 5 and 8).
3. The TEG vaporization loss from top of still/regenerator column is practically independent of the number of theoretical trays in the stripping gas section (N_S), see Figure 4.
4. Pressurized reboiler results in higher TEG vaporization losses from regenerator due to higher stripping gas requirements.
5. Even though not studied in this TOTM, mechanical losses such as entrainment from contactor top and regenerator top as well as leaks from pump seals are much higher than the vaporization losses presented here.
To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), and PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Dr. Mahmood Moshfeghian

References:
1. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2013.
2. Soave, G., Chem. Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.
3. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2013/11/estimating-still-column-top-temperature-in-teg-dehydration-unit/, Tip of the Month, November 2013.
4. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2013/09/high-pressure-regeneration-of-teg-with-stripping-gas/, Tip of the Month, September 2013.
Appendix

Fig 2A. Variation of TEG vaporization loss from regenerator top with circulation volume ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F) and N_S=0

Fig 3A. Variation of TEG loss from regeneartor top with circulation volume ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F) and N_S=0
December 1, 2013
Estimating Still Column Top Temperature in TEG Dehydration Unit
In this Tip of The Month (TOTM), the effect of striping gas rate and TEG circulation ratio on the still column top temperature for regeneration of rich triethylene glycol (TEG) is investigated. Specifically, this study focuses on the variation of still column top temperature with reboiler pressure, TEG circulation ratio and stripping gas rate. By performing a rigorous computer simulation of TEG regeneration at reboiler pressures of 110.3 kPaa (16 psia) and 524.1 kPaa (76 psia), two charts for quick determination of still column top temperature needed for facilities type calculations are developed. In addition, the effect of theoretical number of trays in the stripping gas section is studied.

Computer Simulation Results:

In order to study the impact of stripping gas rate and TEG circulation rate on the still column top temperature, the TEG dehydration process was simulated using ProMax [1] software with its Soave-Redlich-Kwong (SRK) [2] equation of state (EOS). The process flow diagram used for these simulations is shown in Figure 1.

The water-saturated gas with a water content of 915 mg/std m³ (57 lbm/MMSCF) enters the bottom of the contactor column at 37.8°C (100°F) and 6895 kPaa (1000 psia) at a rate of 2.835×10⁶ std m³/d (100 MMSCFD). The contactor column has three theoretical trays. The lean TEG solution enters at the top of the contactor column and flows down in the column. As shown in Figure 1, the water content of the dried gas is 10 mg/std m³ (0.63 lbm/MMSCF). The rich TEG solution contains 96.1 mass percent TEG entering the still column at 100°C (212°F) and 515 kPaa (74.7 psia). The reboiler temperature was set at 204.4°C (400°F) and boil-up ratio of 0.1 (molar bases). Two theoretical trays in the regenerator (still) column (N_R = 2) and two theoretical trays (N_S = 2) in the striping gas section were specified. The striping gas enters the bottom of the stripping gas section at 204°C (399°F) and 524 kPaa (76 psia). Methane was used for the stripping gas at a rate of 56.3 std m³/h (1893 scf/hr). The regenerated lean solution contains 99.6 mass percent TEG and the ratio of stripping gas to lean TEG liquid volume rates is 20 std m³ of gas/std m³ of lean TEG solution (2.67 scf/sgal) or a mass ratio of 28.3. The regenerator (still) top temperature is 91.4°C (196.5°F). If the same stripping gas was sparged directly into the reboiler (N_S= 0, no stripping gas section), with everything else remaining the same, the regenerated solution contains 99.2 mass percent TEG and the regenerator column top temperature remains practically the same and is 91.1°C (196°F). For the above case the number of theoretical trays in the still column is increased from 2 to 3 (N_R = 3); the lean TEG concentration increased slightly from 99.6 to 99.8 mass percent but the regenerator column top temperature remained the same.

Using a similar set up as is shown in Figure 1, several simulations were performed for a range of stripping gas rates, for N_R=2, N_S=0 and for two reboiler pressures of 110.3 and 524 kPaa (16 and 76 psia) and temperature of 204.4°C (400°F). The results of these simulation runs are presented in Figures 2 to 5.

Figure 1. Sample results using ProMax [1] for TEG dehydration with reboiler P=110.3 kPaa (16 psia) with N_R=2 and N_S=2

Figures 2 presents the variation of still column top temperature with circulation ratio (mass basis) and stripping gas rate at top pressure of 101.3 kPaa (14.7 psia) and reboiler pressure of 110.3 kPaa (16 psia) operating at 204.4°C (400°F).

As was discussed in the August 2013 TOTM, regeneration of TEG at higher reboiler pressure has several advantages such as preventing the emission of harmful contaminants like benzene, toluene, ethylbenzene, xylenes (BTEX), and hydrogen sulfide to the environment [3]. Therefore, similar diagrams as shown in Figure 2 were generated for top pressure of 515.2 kPaa (74.7 psia) and reboiler pressure of 524.1 kPaa (76 psia) at 204.4°C (400°F). Figure 3 presents the variation of still column top temperature for such a high reboiler pressure.

Fig 2. Variation of still column top temperature with circulation mass ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

Fig 3. Variation of still column top temperature with circulation msss ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

Figures 2 and 3 can be used for a quick determination of the still column top temperature for a given stripping gas rate and TEG circulation ratio either at low or high reboiler pressure. The two reboiler pressures selected in this study are typical operating pressures. For generation of data for Figures 2 and 3, the stripping gas was sparged directly into the reboiler; therefore, the number of theoretical trays for stripping gas section is zero (N_S=0). The corresponding figures in terms of TEG circulation volume ratio are presented in the Appendix (Figures 2A and 3A).

Generally, either 0, 1, or 2 theoretical trays in the stripping gas section is used. In order to investigate the effect of the number of theoretical trays in the stripping gas section (N_S) on the still column top temperature, simulations were performed for the cases of N_S=0 and N_S=2 for two constant stripping gas rates.

Figures 4 and 5 present the results of these simulations for low and high reboiler presssures of 110.3 kPaa (16 psia) and 524.1 kPaa (76 psia), respectively. The reboiler temperature for all cases was set at 204.4°C (400°F).

Figures 4 and 5 clearly indicate that the still column top temperature is independent of the number of theoretical trays in the stripping gas section. Therefore, Figures 2 and 3 can be used for any number of theoretical trays in the stripping gas section.

Fig 4. Effect of the number of theoretical trays (N_S) on the still column top temperature at various circulation ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

Fig 5. Effect of the number of theoretical trays (N_S) on the still column top temperature at various circulation ratio and stripping gas rates at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

Similar study also showed that the feed gas temperature to the contactor column has no effect on the still column top temperature. The results of this study are shown in Figures 6 and 7 of the Appendix.

Conclusions:

In this TOTM, the effect of circulation ratio, stripping gas rate, theoretical number of trays, and the feed gas temperature to the contactor column on the still column top temperature for regeneration of TEG concentration at low and high reboiler pressure operating at 204.4°C (400°F) was studied. Two charts for a quick determination of the still column top temperature at a specified stripping gas rate and circulation ratio to achieve a desired level of lean TEG concentration were prepared and presented in Figures 2 through 3 (see the corresponding figures in the Appendix). These charts are based on the rigorous calculations performed by computer simulations and can be used for facilities type calculations for evaluation and trouble shooting of an operating TEG dehydration unit. In addition, the following observations were made:
1. The still column top temperature is independent of the number of theoretical trays in the stripping gas section (N_S) and feed gas temperature to the contactor column.
2. As the stripping gas rate increased, the still column top temperature decreased.
3. As the TEG circulation ratio increased, the still column top temperature decreased.
4. Pressurized reboiler results in much higher still column top temperature than the atmospheric reboiler.
To learn more, we suggest attending our G40 (Process/Facility Fundamentals), G4 (Gas Conditioning and Processing), G5 (Gas Conditioning and Processing-Special), and PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), courses.

John M. Campbell Consulting (JMCC) offers consulting expertise on this subject and many others. For more information about the services JMCC provides, visit our website at www.jmcampbellconsulting.com, or email us at consulting@jmcampbell.com.

By: Dr. Mahmood Moshfeghian

References:
1. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2013.
2. Soave, G., Chem. Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.
Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2013/08/teg-dehydration-how-does-the-stripping-gas-work-in-lean-teg-regeneration/, Tip of the Month, August 2013.

Appendix

Fig 2A. Variation of still column top temperature with circulation volume ratio and stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

Fig 3A. Variation of still column top temperature with circulation volume ratio and stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)

Fig 6. Variation of still column top temperature with circulation mass ratio and feed gas temperature to the contactor column at a specified stripping gas rate at top P=101.3 kPaa (14.7 psia) and reboiler P=110.3 kPaa (16 psia) at 204.4°C (400°F)

Fig 7. Variation of still column top temperature with circulation mass ratio and feed gas temperature to the contactor column at a specified stripping gas rate at top P=515.2 kPaa (74.7 psia) and reboiler P=524.1 kPaa (76 psia) at 204.4°C (400°F)
November 1, 2013