Campbell Tip of the Month

Category: Gas Processing

  • 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, 9th Edition, 1st Printing, Editors Hubbard, R. and Snow –McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.
    3. GPSA Engineering Data Book, Section 21, Volume 2, 13th 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.

     

  • Surge Control Alternatives for Refrigeration Systems

    Unlike the natural gas compressors which rely upon recycling the after-cooled compressed gas to compressor suction for surge prevention, the surge control in refrigeration systems is quite different because the refrigerant vapors are compressed only to their saturation pressure for dew point temperature, which when cooled for recycling immediately condense into a liquid. Consequently, the hot compressed refrigerant vapors, also known as the hot gas, before any cooling are recycled to compressor suction to make up for any loss of the process chilling load that reduces the amount of available refrigerant vapors to satisfy the minimum flow to the compressor suction, thus avoiding the phenomena known as “surge”.

    Three-Stage Refrigeration Cycle

    A three-stage propane refrigeration system [1, 2] is adapted into Figure 1 to illustrate a typical refrigeration cycle.

    ”]As shown in Figure 1, all three process chillers are located on the 1st stage and there is no process side chilling load at the two economizer levels. The shown schematic is a simplest illustration of a three-stage propane refrigeration cycle [1, 2].

    The compressed refrigerant vapors are condensed with ambient air. The high pressure (D-0116) and low-pressure (D-0115) economizers separate the flashed vapors that enter the third and second stages of the refrigerant compressor, respectively. To protect the compressor from any liquid carry over, each stage has a suction drum, and all vapors entering the compressor must pass through these drums (D-0112, D-0113 and D-0114).

    Surge Protection with Hot Gas and Quench Liquid

    So how do you provide a cool vapor stream to prevent the refrigerant compressor from surging in order to keep the discharge temperature from getting too high? To address this challenge, hot compressed vapors from refrigerant compressor discharge are used to vaporize liquid refrigerant, in a similar manner to how the process stream is used to vaporize refrigerant in the chillers.

    In the refrigeration cycle of Figure 1, the hot gas from compressor discharge flows into each suction drum as determined by the anti-surge controller to prevent each compression stage from surging. To cool the hot recycled vapors, a temperature controller introduces quench liquid directly from the refrigerant accumulator D‑0111 into the inlet of each compressor suction drum. Ideally, if the compressor anti-surge vapor flow and temperature control valve on quench liquid are perfectly matched and there is no time lag within the control system, the configuration of Figure 1 will work perfectly.

    Unfortunately, if there is any mismatch, which in a real world situation always exists,   either too much quench liquid is added to the inlet stream whereby there is liquid level in the suction drums or not enough quench liquid is added whereby the discharge temperature keeps rising. In either mismatch case, the refrigerant compressor will shut down either from high liquid level in the suction drum or from high discharge temperature caused by the high inlet temperature due to recycle of hot gas. 

    Surge Protection with Sparged Hot Gas using No Quench Liquid

    A well-proven alternative system is shown in Figure 2 [3, 4] in which hot propane vapors from the compressor discharge, as controlled by the anti-surge system, flow through the sparged line submerged in the liquid filled portions of the HP (D-0116) and LP (D-0115) economizers, and under the tube bundle of the first stage chillers (E-0107, E-0108, and E-0111). There is no need for quench liquid as the hot vapors provide the heat to the available liquid propane in these services to generate the required propane vapors to prevent the compressor surge conditions.

    ”]The configuration of Figure 2 builds upon the recognition that:

    • The shortage of propane refrigerant entering the compressor stages is caused by reduction in heat exchanger duty from changes on the process side of the chillers (E-0107, E-0108 or E-0111). This reduction of heat transfer duty is offset by routing the required amount of hot propane vapor from compressor discharge directly to the kettle side of the chiller to supplement process thermal requirements.
    • The compressor suction drums are provided to prevent any carried-over liquid from entering the compressor and should essentially have no liquid in them.
    • By diverting the anti-surge hot vapor away from the compressor suction drums assures that the compressor suction drums stay dry.
    • By using the already present liquid refrigerant in the chillers and economizers it avoids any potential mismatch between the hot propane vapors and the quench liquid that may inadvertently buildup high liquid level and end up tripping the refrigerant compressor.

    Conclusion

    While both refrigerant compressor surge control systems are proven and used by the industry, the hot sparged gas with no quench liquid configuration of Figure 2 is more forgiving and reliable.

    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: Yuv R. Mehra

    References:

    1. Al-Shahrani, Saleh M. and Mehra, Yuv R., “Start-up Experience for Value Recovery from CCR Net Gas at Yanbu’ Refinery,” Advances in Hydroprocessing II Session, Paper 26d, 2007 Spring National Meeting of AIChE, Houston, Texas, USA, April 22-26, 2007.
    2. Al-Shahrani, Saleh M. and Mehra, Yuv R., “Saudi Aramco installs new LPG recovery unit at Yanbu’ Refinery,” Oil & Gas Journal, Vol. 105.21, Jun. 4, 2007, p. 60.
    3. Mehra, Yuv R., “Focus on Value-Chain Contributions during Process Technology Selection,” Worldwide Developments Forum, 86th Annual Meeting of the Gas Processors Association, San Antonio, Texas, USA, March 11-14, 2007.
    4. Mehra, Yuv R., “Focus on Value-Chain Contributions during Process Technology Selection,” Saudi Aramco Journal of Technology, Summer 2007, p. 2.

  • 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”, 9th 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.

  • 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 lbm/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.

  • 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, 13th 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.

  • Debriefing Jobs Provides Several Benefits Associated With Process Safety

    A pillar of Risk Based Process Safety (RBPS) is Learn from Experience.  The work we do and the processes we use to analyze our work provide significant learning opportunities to enhance process safety competency.  This is a derivative of Kolb’s experiential learning cycle [1], but many times we fail to take advantage of the learning opportunities available to us unless there is an incident or a near miss.

    This Tip of the Month (TOTM) will introduce a simple method for debriefing the job tasks we perform to close the loop on this cycle and capture appropriate data to develop competency, work safely and capture near miss/incident data quickly and efficiently.

    Conducting a simplified job hazard analysis will ensure that all hazards are identified, managed, and mitigated prior to performing work.  Performing a simple debrief at the conclusion of the work will ensure that we learn from the experience. By considering every job to be performed a learning opportunity, the experiential learning cycle can be used to identify what was done, how well it was done, and how we might improve in the future.

    This Month’s Tip was recently presented at the Mary K. O’Connor Process Safety Symposium at Texas A&M University [1].

    One of the pillars of the Center for Chemical Process Safety’s (CCPS) Guidelines for Risk Based Process Safety is “Learn from Experience.”  What does this mean?

    The elements of this pillar include:

    • auditing,
    • management review and continuous improvement,
    • measurement and metrics and
    • incident investigation.

    Each of these elements provides findings, lessons and data that are useful for learning and thus changing and enhancing behaviors and attitudes.  The change and enhancement will influence an organization’s culture and ultimately push the organization toward a learning culture.

    These are not the only opportunities available for organizations to learn from experience.  Metrics and audits can allow a general overview of process safety performance.  Incident investigation insures that when reported, incident information is transmitted to all who will benefit from the learning.

    The job hazard analysis process that many organizations use to identify and mitigate hazards provides a tremendous opportunity to capture data and use the experiential learning cycle if the job is debriefed properly after completion.  This paper will provide guidance and explain the benefits that can be derived from debriefing completed jobs.

    At the 2008 symposium, this author presented a paper entitled “Three Simple Things to Improve Process Safety Management.”  One of those simple things was to conduct a formalized Job Hazard Analysis (JHA) for the tasks being performed in the life cycle of a process.  That paper presented a checklist that could be used to guide personnel in the process of conducting a JHA.  (See checklist at end of this paper)

    Many facilities have embraced the concept of conducting JHA.  They may be called something else  (job safety analysis, job safety checklist, job task analysis) but the process is essentially the same.  The job or task is identified and analyzed step by step.  The analysis is to identify hazards that may be involved with each step and then develop strategies to mitigate the hazards.  This sounds simple in theory, but in reality there are many things that can and do go wrong with this process.

    To provide consistency and to make it easier to track that these analyses have been completed, standardized checklists and forms have been created that list the most common hazards that can be found with a job and logically guide the user toward identification and mitigation of hazards.  Experience shows that after these forms and checklists have been used regularly, some personnel have a tendency to try and short cut the process.  This leads to what is known as “pencil whipping” the JHA.  In other words, personnel will complete the checklist or form without actually performing the analysis required.   Familiarity with the forms and checklists may drive personnel to identify common hazards, but do little to mitigate the hazards.  For example, a common checklist item is “slips, trips and fall hazards”.  Personnel will identify that the ground is rutted or that there is ice on the ground, but few will actually smooth the ground or cover the ice with sand to mitigate the hazards identified.

    It is generally agreed among those who supervise personnel performing JHAs that the most important part of the process is not the completion of the forms and checklists but the discussion that happens among a group performing the work.  In order to focus the discussion and insure that all issues are addressed, the JHA checklist at the end of this paper can be used.  The JHA checklist is not intended to replace the checklists and forms that an organization may already have in place.  The JHA checklist can enhance the process by focusing a group’s thoughts on individual checklist items.  By answering each question a work group should be able to identify all issues associated with any job they are conducting.

    As work groups become more familiar with the JHA checklist and the process of discussing and documenting the efforts of the group, a simplified method can be adopted.  By answering six key questions, a group of workers can focus discussion on the issues that are most important.   The six questions and the benefits of using them include:

    What are we doing?  If we can’t answer this question completely and in simple terms, then we should not be doing the job.  A simple explanation will insure that all members of the team are working toward the same goal.

    What is the most dangerous part?  If we can identify the most dangerous part of what we are doing we have identified all hazards, ranked them and determined the most dangerous part.

    What will we do to protect ourselves?  Answering this question ensures that all mitigation measures have been put into place and that all personnel know what is being done.

    How will we know we are changing what we are doing?  To answer this question effectively, we will need to be creative and analytical.  Examination of the work site, knowledge of simultaneous operations, and competency in our job will be required.  Anticipating potential changes will insure that we are not surprised when things do change.

    What will we do about it?  Again, creativity and analytical thinking are critical here.  Many times we hear the phrase, “prior planning prevents poor performance”.  Effectively answering this question insures that performance will be as designed.

    How will we know we are finished?  Review of completed job hazard analysis documents has shown that it may be difficult to determine at what point the job is complete.  If the permit for the job being performed provides a scope of work like, “replace mechanical seal in hot oil pump”, once the seal is replaced, there are numerous tasks that still need to be performed before the job is complete.  Numerous times the JHA does not go beyond analyzing the tasks associated with the scope of work and do not consider additional tasks; like testing, clean up and turnover to operations.

    As previously mentioned, most supervisors believe that the discussion associated with this type of analysis is more important than the completion of the form used to show that the JHA has been performed.

    What about the form though?

    • What happens at the conclusion of the job?
    • Does anyone review the form to determine if all the hazards were found and mitigated?
    • Does anyone follow up with the work group to see if anything happened that made them change the work?
    • How should this review be performed and what are the benefits that will be gained by this?
    • How can we learn from our experience?

    Developing competent personnel is an ongoing process for most organizations.  A great deal of literature exists on the most effective methods of developing competency in adults. Training sessions are delivered using the concept of Kolb’s theory of the experiential learning cycle.  According to Kolb [2], this type of learning can be defined as “the process whereby knowledge is created through the transformation of experience.” [i] In other words, adults learn best when they are actively experiencing something and not just listening to lectures or instructor centered learning.

    Experienced trainers who deliver adult learning sessions use a process of debriefing to allow reflection, reinforce learning and help the learner apply the knowledge to their life.  It is generally acknowledged in the training industry that most real learning takes place in the debrief.  This is the opportunity for learners to reflect and develop knowledge from the activity, in our case the job performed.

    Very simply, debriefing a learning activity should focus on three questions.  What?  So What?  Now What?

    What? is the question that guides the learning toward reflection and what just happened.  This question provides a starting point to discover what everyone involved experienced.

    So What? is the question that leads to drawing conclusions and exploring alternate methods.

    Now What? leads to future planning and continuous improvement initiatives that will be used to strengthen the organization’s culture and work processes.

    If we return to question six of the job hazard analysis process, “How will we know we are done?”, the final answer for this question would be, “When we have completed the debrief of the job performed.”  There are five questions that should be used for debriefing a job.  These five questions, how they relate to the standard debriefing questions and the expected lessons to learn from them include:

    What did we do?  This is the opportunity for reflection and to insure that the job has been completed appropriately.  Each member of the team should come to agreement that what is being described is what was actually done.   This is the What of debriefing.

    Did anything change while doing the job?   Reflection on this question will lead the team to determine if the job was actually performed as it was initially described and analyzed.  This is the question that will also lead to identify incidents for investigation.  If anything unusual occurred during the task, reporting should be more efficient because the incident will be fresh in everyone’s mind.  Capturing these incidents and changes now will help modify future work orders and insure that we learn something from this experience.  This is the So What of the debriefing cycle.

    Did anybody get hurt?  This question should be answered with all personnel examining themselves for strains, pulled muscles, bumps, bruises, cuts, scrapes, twisted joints, twinges in the back and a general self examination for good health.  Any small injury or potential illness should be recorded here.   This will insure that a worker does not leave the job without reporting an injury or illness, and then visit a medical provider later because something cropped up.  Having someone discover they have been injured after leaving the worksite is a problem for managers.  This allows measures to be taken early to manage the injury or illness for reporting purposes.  Here and the next question is where more exploration of the “What” is performed.

    Did anybody come close to getting hurt?  This is the question that will capture near miss incidents quickly.  Near miss reporting programs fail for numerous reasons.  Lack of understanding, lack of motivation, blaming the reporter, and convenience of reporting are reasons that near misses may not be reported.  Reflection and discussion about the completed job will insure that any near miss is reported quickly.  This will lead to creation of a more comprehensive database that can be used to predict trends and identify problems areas in processes.

    What would we do differently?  This is the question that will tie everything together into a plan for the future.  Recommendations and action items should be generated from this final question so that future jobs can be analyzed with more speed and efficiency.  Potential training and competency development issues may be discovered.  Procedures for modification may be identified.  Latent conditions that are not readily apparent may be identified and mitigated before they become active failures.

    The Now What of the debriefing cycle is:

    • Conducting an effective job task analysis and following with an effective debriefing of the job will yield several benefits.
    • Competency gaps of personnel associated with the work will be identified.
    • Training topics and on the job mentoring for personnel with these identified gaps, can be more quickly delivered.
    • Procedural modifications that are necessary to insure that work is performed safely and efficiently will be quickly identified and addressed.
    • Potential process safety incidents will be quickly identified and investigated.
    • Near miss incidents will be reported quickly and the organization’s near miss/incident database will be enhanced.

    The process described in this paper can be expanded to any job and any work group.  Consider an engineering team who is working on the design of a new process to be considered for construction.  Conducting an effective job task analysis in the beginning stages of the project will insure that roles, goals and expectations are addressed and known.  Conducting an effective debrief at the conclusion, or even at selected stages of a project, will enhance the project team’s effectiveness and insure that all team members are always striving to meet the goal of the project.  The attached checklist for engineering projects, at the end of this paper, may be helpful for focusing a team’s efforts.

    Opportunities exist in all phases of operations and in all activities performed to keep processes safe.  It is important that all personnel be aware that learning from experience happens every day and these lessons learned need to be captured and stored for future use.

    To develop process safety competency attend our PS-4, Process Safety Engineering;  HS-45, Risk Based Process Safety Management; and PS-2, Fundamental of Process Safety courses.  To develop competency in other skills, attend one of our other courses.

    By Clyde Young

    PetroSkills Instructor/Consultant

    Reference:

    1.    Young, Clyde. ,” Debrief:  The experiential learning cycle, process safety competency, safe work practices, identifying and reporting of near miss/incident data”, Mary K. O’Connor Process Safety Symposium, Texas A&M University, October 29.

    2.    Kolb, David A. Experiential Learning: Experience as the Source of Learning and Development. Prentice-Hall, Inc., Englewood Cliffs, N.J. 1984.

    Job Hazard Analysis Checklist

    1. PROCEDURES

    • ·What are the procedures for the task?
    • ·What is unclear about the procedures?
    • ·What order will we use these procedures?
    • ·What permits are needed for hazard controls?

    2. EQUIPMENT AND TOOLS

    • ·What are the right tools for the job?
    • ·What is the correct way to use them?
    • ·What is the condition of the tool?

    3. POSITIONS OF PEOPLE

    • ·What could we be struck by?
    • ·What could we strike ourselves against?
    • ·What can we get caught in/on/between?
    • ·What are potential trip/fall hazards?
    • ·What are potential hand/finger pinch points?
    • ·What extreme temperatures will we be in/around?
    • ·What are the risks of inhaling, absorbing, swallowing hazardous substances?
    • ·What are the noise levels?
    • ·What electrical current/energized system could we come in contact with?
    • ·What would be a cause for overexerting ourselves?

    4. PERSONAL PROTECTIVE EQUIPMENT

    • ·What is the proper PPE?

    Hard hat, glasses/goggles, ear plugs, gloves, steel toe boots, respiratory system, fire retardant clothing

    5. CHANGING THE COURSE OF WORK

    • ·What would cause us to have to stop or rearrange the job?
    • ·What would cause us to change our tools or equipment?
    • ·What would cause us to have to change our position?
    • ·What would cause us to have to change our PPE?

    YOU HAVE THE RIGHT AND

    THE OBLIGATION TO

    STOP UNSAFE ACTS

    ENGINEERING JOB ANALYSIS

    1. PROCEDURES

    • ·What are the procedures for the task?
    • ·What is unclear about the procedures?
    • ·In what order will we use these procedures?
    • ·What is the proper timeline for these procedures?
    • ·What permits or permissions are needed for job controls?

    2. EQUIPMENT, TOOLS, DOCUMENTS

    • ·What are the right tools for the job? (software, simulators, matrixes, checklists, worksheets…)
    • ·What is the correct way to use them?
    • ·What forms will be needed for the job?
    • ·What documents will we need to produce?
    • ·What else do we need to know?

    3. INTERACTION WITH PEOPLE

    • ·What other departments need to know about this task?
    • ·Who are the personnel that need to know?
    • ·What other departments will supply information for this task?
    • ·Who are the personnel who will supply that information?
    • ·What could prevent other personnel or departments from supplying what we need?
    • ·What could prevent us from supplying what other departments need?

    4.  CHANGING THE COURSE OF WORK

    • ·What would cause us to have to stop or rearrange the job?
    • ·What would cause us to change our tools or equipment?
    • ·What would cause us to have to change our interaction with people?

    YOU HAVE THE RIGHT AND THE OBLIGATION TO

    STOP UNSAFE or UNPRODUCTIVE ACTS

  • 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 (NR = 2) and two theoretical trays (NS = 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 (NS = 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 (NR = 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 NR=2, NS=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 NR=2 and NS=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 NS=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 NS=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 (NS=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 (NS) on the TEG vaporization loss, simulations were performed for the cases of NS=0 and NS=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 (NS) 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), NS=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 NS=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 NS=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) , NS=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 (NS), 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 NS=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 NS=0

  • 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 (NR = 2) and two theoretical trays (NS = 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 (NS = 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 (NR = 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 NR=2, NS=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 NR=2 and NS=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 (NS=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 (NS) on the still column top temperature, simulations were performed for the cases of NS=0 and NS=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 (NS) 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 (NS) 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 (NS) 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)

  • High Pressure Regeneration of TEG with Stripping Gas

    In this Tip of The Month (TOTM), regeneration of rich triethylene glycol (TEG) with striping gas at high pressure is investigated. Specifically, this study focusses on the determination of the required stripping gas rate as a function of the lean TEG mass percent, reboiler temperature, and the number of theoretical trays in the stripping section (NS) for a regenerator (still) column with two theoretical trays (NR). By performing rigorous computer simulations of TEG regeneration at high pressure, a series of charts for quick determination of stripping gas rates needed for facilities type calculations are developed. The results of this study are also compared with the case of TEG regeneration at atmospheric pressure, which was published in the June 2013 TOTM.

    In gas dehydration service, TEG absorbs limited quantities of benzene, toluene, ethylbenzene, and xylene ( BTEX) along with other volatile organic compounds (VOCs)  from the gas. BTEX and other VOC emissions are an environmental challenge for the natural gas industry since some of these compounds are considered to be carcinogenic.

    Absorption is fa­vored at lower temperatures, higher pressure, increased lean TEG concentration and circulation rate. Based on the data from the June 2011 TOTM [1], predicted absorption levels for BTEX components vary from 5 to 15% for benzene, 5 to 25% for toluene and ethylbenzene, and 5 to 35% for xylene [2]. Some of these absorbed BTEX components are released with the flashed gas off of the TEG separator operating at around 483 kPag (70 psig). The flashed gas may be used as fuel or sent to the flare. The rich TEG solution is normally regenerated at low pressure and high temperature. The remaining BETX components are released with the vaporized water and stripping gas at the top of the still column. Hicks et al. [3] discuss different options for VOC emission control. One of their options was to operate the reboiler at pressures that allow the overhead vapors that contain VOCs including BTEX and stripping gas to flow directly to either compressor suction, a fuel system, or to a flare. Their study showed higher pressure reboilers effectively control BTEX emissions and economically recover all the hydrocarbon vapors with minimal incremental capital cost and no required emissions monitoring. Hicks et al. [3] highlighted the following advantages for the pressurized glycol regeneration system:

    • All vapors (VOCs, H2S, CO2) released from the glycol are at a sufficient pressure, which allows the use of simple handling methods without discharge to the atmosphere. Typical handling methods are:
      1. Mixing the released vapors with fuel gas or other gas users.
      2. Using compressors in other service instead of dedicated compression to recompress the released vapors.
      3. Using dedicated compression which allows for better reboiler pressure operation.
    • Because the previously mentioned methods can handle vapors from the still overheads, one can eliminate the equipment specifically designed to recover these vapors. This equipment typically includes condensers, separators, pumps, and a means for burning or incinerating the non-condensable vapors. The entire still column overhead can be routed to the compression suction scrubbers with no intervening components like a condenser or separator.
    • The larger-than-normal amount of stripping gas used in the stripper results in less VOCs in the condensed water vapor from the overheads. This is because the vapor-liquid equilibrium is shifted such that condensation of VOCs is minimized. The disposal water will therefore contain significantly less undesirable components. Similarly, for those systems in which the overheads are mixed with other gases, the volume ratio of overhead to mixing gas (such as gases for fuel and compressor suction) will further reduce the VOCs in the condensed water.
    • No environmental testing for atmospheric discharge of VOCs is required because no VOCs will be released, as is often the case with conventional technology. Furthermore, discharge of VOCs from the regeneration system at considerably higher-than-normal pressures facilitates mixing and extreme dilution of VOCs in plant fuel or in fuel sold and transmitted to remote users.

    Computer Simulation Results: 

    In order to study the impact of stripping gas rate on the lean TEG mass percent, the TEG regeneration process was simulated using ProMax [4] software with its Soave-Redlich-Kwong (SRK) [5] equation of state (EOS). The process flow diagram used for these simulations is shown in  Figure 1.

    Figure 1. Sample results using ProMax [3] for high pressure TEG regeneration at reboiler P=515 kPaa (74.7 psia) and NS=2

    As shown in Figure 1, the rich TEG solution contained 97.5 mass percent TEG entering the still column at 150°C (302°F) and 515 kPaa (74.7 psia). The reboiler temperature was set at 204°C (400°F) and boil-up ratio of 0.1 (molar bases).  Two theoretical trays in the still column (NR = 2) and two theoretical trays (NS = 2) in the gas striping section were specified. The striping gas enters the bottom of the gas stripping section at 150°C (302°F) and 521.9 kPaa (75.7 psia). Methane was used for the stripping gas at a rate of 70 std m3/h (2458 scf/hr). The regenerated lean solution contains 99.25 mass percent TEG and the ratio of stripping gas to lean TEG liquid volume rates is 80.26 std m3 of gas/std m3 of lean TEG solution (10.73 scf/sgal). If stripping gas was sparged directly into the reboiler (NS = 0, no gas stripping section), with everything else remaining the same, the  regenerated solution contains 98.49 mass percent TEG and  the ratio of stripping gas to lean TEG liquid volume rates is 79.58 std m3 of gas/std m3 of lean TEG solution (10.64 scf/sgal). For the above cases the number of theoretical trays in the still column is increased from 2 to 3 (NR = 3) and the lean TEG concentration remained almost the same. The concentration of rich TEG solution is also varied from 90 to 98 mass percent, but no appreciable change in the lean TEG concentration was observed for the same stripping gas rate.

    Using a similar set up as is shown in Figure 1, several simulations were performed for; a range of stripping gas rates, for NR=2, NS=0, 1, 2 and for reboiler temperatures of 204, 193, and 182°C (400, 380, and 360°F) and a reboiler pressure of 515 kPaa (74.7 psig). The results of these simulation runs are presented in Figures 2 to 4. For purposes of comparison all of these diagrams are replotted in Figure 5.

    Fig 2. Effect of lean TEG mass %, reboiler temperature and number of theoretical trays in stripping gas column (NS=0) at reboiler P=515 kPaa (74.7 psia) on the stripping gas rate

    Fig 3. Effect of lean TEG mass %, reboiler temperature and number of theoretical trays in stripping gas column (NS=1) at reboiler P=515 kPaa (74.7 psia)  on the stripping gas rate

    Fig 4. Effect of lean TEG mass %, reboiler temperature and number of theoretical trays in stripping gas column (NS=2) at reboiler P=515 kPaa (74.7 psia)  on the stripping gas rate

    Figure 6 presents the required stripping gas rate for both a pressurized reboiler, 515 kPaa (74.7 psia), and a low pressure, 109 kPaa (15.8 psia) reboiler having two theoretical trays (NS=2) in the gas stripping section. This figure indicates that required stripping gas rates are 10 to 100 times higher for pressurized reboiler when compared to the low pressure reboiler.

    Conclusions:

    In this TOTM, the effect of stripping gas rate on the regenerated lean TEG concentration at high pressure for several operation conditions was studied. A series of charts to use for a quick determination of the required stripping gas rate to achieve a desired level of lean TEG concentration are presented in Figures 2 through 6. 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 an operating TEG dehydration unit. In addition, the following observations were made:

    1. For rich TEG concentrations between 90 and 98 mass percent, the required stripping gas rate is independent of rich TEG concentration.
    2. As the number of theoretical trays in the stripping column (NS) increases from 0 to 2, the required striping gas rate decreases.
    3. Increasing the number of theoretical trays in the still column (NR) from 2 to 3 has no appreciable effect on the stripping gas requirement.
    4. Increasing the reboiler temperature from 182 to 204 °C (360 to 400 °F), decreases the required stripping gas rate.
    5. Increasing the TEG reboiler pressure from 109 kPaa (15.8 psia) to 515 kPaa (74.7 psia) increases the stripping gas requirement by a factor of between 10 and 100 depending on other factors.

    Fig 5. Effect of lean TEG mass %, reboiler temperature and number of ideal trays in stripping gas column on the stripping gas requirement at reboiler P=515 kPaa (74.7 psia).

    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

    Fig 6. Comparison of high pressure with low pressure stripping gas requirements for TEG regeneration. HP=515 kPaa (74.7 psia), LP=109 kPaa (15.8 psia) with NS=2

    References:

    1. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2011/06/absorption-of-aromatics-compounds-in-teg-dehydration-process/, Tip of the Month, June 2011.
    2. Moshfeghian, M. and R.A. Hubbard, “Quick Estimation of Absorption of Aromatics Compounds (BTEX) in TEG Dehydration Process,” Proceedngs of the 3rd International Gas Processing Symposium, March 5 – 7 2012 , Qatar , 2012.
    3. Hicks, R., Gallaher, D. and R. Craig, “Pressurized reboiler reduces VOC emissions in glycol dehy systems”, Oil & gas j., Vol 102, Issue 17, April 2004.
    4. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2013.
    5. Soave, G., Chem. Eng. Sci. Vol. 27, No. 6, p. 1197, 1972.

  • TEG Dehydration: How Does the Stripping Gas Work in Lean TEG Regeneration?

    For dehydration of natural gas by triethylene glycol (TEG) process to a lower water content/water dew point temperature a higher lean TEG concentration is required. To achieve a higher lean TEG concentration at a specified reboiler temperature and pressure, commonly a stripping gas is used. Stripping is defined as a physical separation process by which one or more components are removed from a liquid stream by a vapor stream. Stripping gas lowers the water partial pressure causing more water to vaporize from TEG solution. Normally, stripping is performed at the practical/possible highest temperature and lowest pressure. This allows for minimum stripping gas flow rate. Any inert gas or a portion of the gas being dehydrated is suitable.

    In this Tip Of The Month (TOTM), the impact of stripping gas on the water partial pressure and lean TEG solution boiling temperature or lean TEG concentration is studied. Two case studies are considered. In case one, at a specified reboiler temperature and pressure, the impact of stripping gas on lowering the water partial pressure and consequently increasing the TEG concentration in the lean TEG solution is demonstrated. Similarly in case two, for a specified 99 and 99.4 mass percent TEG in the liquid phase and atmospheric pressure, the impact of stripping gas on lowering the water partial pressure and consequently lowering the liquid boiling point (bubble point) temperature is demonstrated.

    Glycol dehydration is the most common dehydration process used to meet pipeline sales specifications and field requirements (gas lift, fuel, etc.). TEG is the most common glycol used in these absorption systems. At atmospheric pressure and a maximum reboiler temperature of 204 °C [400 °F] the highest glycol concentration of lean TEG that can be achieved is roughly 99 mass percent. This represents the maximum lean TEG concentration that can be produced in a reboiler operating at 1 atm. If the lean TEG concentration required at the absorber to meet the water dew point specification is higher than 99 mass percent, then some method of further increasing the glycol concentration at the regenerator must be incorporated in the unit. Virtually all of these methods involve lowering the partial pressure of water in the glycol solution either by pulling a vacuum on the regenerator or by introducing stripping gas into the regenerator [1].

    In the June and July 2013 Tip Of The Months (TOTM) [2, 3], the effect of stripping gas rate on the regenerated lean TEG concentration for several operation conditions was studied. A series of charts and their corresponding correlations for quick determination of the required stripping gas rate to achieve a desired level of lean TEG concentration were presented. The charts were based on the rigorous calculations performed by computer simulations with correlations developed by regressing the same data to a proposed model. These can be used for facilities type calculations for evaluation and trouble shooting of an operating TEG dehydration unit. For further detail refer to the June and July 2013 TOTM [2, 3].

    Vapor-Liquid Equilibrium:

    For the sake of simplicity, pure methane is used as the stripping gas; however, the following discussion is valid for any gas mixture.

    The governing equations for vapor-liquid equilibrium of TEG, water, and methane are derived and presented in Appendix A at the end of this TOTM. Even though the system of water, TEG and methane is non-ideal, to show the principle of gas stripping, assume the vapor phase is an ideal gas and the liquid phase is an ideal solution. Under these conditions Raoult’s Law for water (See Eq A5 in Appendix A) is.

    According to Raoult’s Law, at a fixed pressure and temperature (i.e. the reboiler condition), the mole fraction of water in the liquid phase (xwater) decreases as water partial pressure decreases. Note that at a fixed temperature, water vapor pressure is constant (see Figure 1). In addition, system pressure (P) is the sum of partial pressure of all components. Assuming an ideal mixture, mathematically P is approximated by:

    Figure 1 presents the vapor pressure of water [4] and TEG [5] for an operating range of reboiler temperature. As can be seen in Figure 1, the vapor pressure (volatility) of water is much higher than that of TEG. In general the vapor pressure of water is approximately 100 to 1000 times greater than that of TEG. This shows why water is vaporized much more than TEG from a TEG solution by heating.

    Since TEG volatility is much lower than that of water and methane, its mole fraction in the vapor phase will be much smaller and for simplicity its partial pressure can be ignored. Therefore equation 2 reduces to:

    At a fixed system pressure (reboiler condition), equation 3 clearly shows that as the methane partial pressure increases, water partial pressure has to decrease and consequently, based on Raoult’s Law, equation 1, the mole fraction of water in the liquid phase decreases. In other words, increasing the stripping gas rate results in lower water mole fraction in TEG and higher TEG mole (mass) fraction is achieved.

    Fig 1A (SI). Comparison of water [4] and TEG [5] vapor pressure

    Fig 1B (FPS). Comparison of water [4] and TEG [5] vapor pressure

    The effect of stripping gas can be summarized as follows:

    • Higher stripping gas (methane) rate increases methane mole fraction in the vapor phase
    • Higher methane mole fraction in the vapor phase increases partial pressure of methane
    • Higher partial pressure of methane lowers partial pressure of water (P is fixed, Eq 3)
    • Lower partial pressure of water lowers mole fraction of water (Raoult’s Law, Eq 1) in the liquid phase (Raoult’s Law, Eq 1) which, in turn, increases the TEG concentration

    In order to demonstrate quantitatively the impact of stripping gas on water partial pressure and lean TEG concentration or boiling temperature and to support the above reasoning, two case studies were investigated. For both cases, ProMax [6] with the SRK (Soave-Redlich-Kwong) [7] equation of state option was used to perform the calculations. In addition, all of the results presented are are for a single equilibrium (theoretical) stage.

    Case Study 1:

    In this case, the impact of stripping gas on regeneration of lean TEG concentration at reboiler conditions of 101.3 kPa and 199°C (14.7 psia & 390°F) is studied.

    Figures 2 A & B present the variation of partial pressures of TEG, water and methane as a function of stripping gas (methane) rate and concentration, respectively. The total pressure is constant and plotted in this figure as well. These figures clearly show that as stripping gas (methane) rate or concentration increases, the methane partial pressure increases and the water partial pressure decreases, approximately in the same amount. Note the TEG partial pressure is relatively constant and very low compared to methane and water. It makes up less than 7 percent of the total pressure.

    Figures 3 A & B present the impact of stripping gas rate and concentration in the vapor and liquid phases on lean TEG concentrations. As can be seen in these figures, as the stripping gas rate (or concentration) increases, the lean TEG concentration increases. It should be emphasized that these results are for a single equilibrium stage. As shown in the previous TOTM, increasing the number of stripping stages changes this relationship significantly.

    Figures 2 and 3 indicate that variation of methane and water partial pressures or lean TEG concentration with stripping gas (methane) rate is non-linear whereas their variation with stripping gas concentration is linear.

    Fig 2A. Impact of stripping gas rate on partial pressure of methane, water and TEG at 101.3 kPa & 199°C (14.7 psia & 390°F)

    Fig 2B. Impact of stripping gas concentration on partial pressure of methane, water and TEG at 101.3 kPa & 199°C (14.7 psia & 390°F)

    Fig 3A. Impact of stripping gas rate on lean TEG concentration at 101.3 kPa & 199°C (14.7 psia & 390°F)

    Fig 3B. Impact of stripping gas concentration on lean TEG concentration at 101.3 kPa & 199°C (14.7 psia & 390°F)

    Case Study 2:

    This is an academic case and is presented to reemphasize the impact of stripping gas. In this case, the impact of stripping gas concentration on the lean TEG solution boiling (bubble point) temperature for relatively fixed lean TEG concentration of TEG and water at 101.3 kPa (14.7 psia) is studied.

    Figure 4 indicates that as the stripping gas (methane) concentration increases, the lean TEG solution boiling (bubble point) temperature drops.

    Figure 5 presents a variation of partial pressures of TEG, water and methane as a function of the stripping gas (methane) concentration. The total pressure is constant and plotted in this figure, too. As in case 1, this figure also clearly indicates that as the stripping gas (methane) concentration increases, while methane partial pressure increases water partial pressure decreases, approximately in the same amount. Note that TEG partial pressure decreases due to the lowering of boiling temperature and its value is much lower compared to those of methane and water. Again, it makes up less than 7 percent of total pressure.

    Variation of component K-values as a function of TEG solution boiling (bubble point) temperature is presented in Figure 6. This figure indicates that the relative volatility of water with respect to TEG (Kwater/KTEG) is in the order of 100 and that of methane with respect to TEG (Kmethane/KTEG) is in the order of more than 10,000.

    Fig 4. Impact of stripping gas concentration on the lean TEG solution boiling (bubble point) temperature at 101.3 kPa (14.7 psia).

    Fig 5. Impact of stripping gas concentration on the component partial pressures at 101.3 kPa (14.7 psia) for 99 mass % TEG and 1 mass % water.

    Fig 6. Variation of K-values with the lean TEG solution boiling (bubble point) temperature at 101.3 kPa (14.7 psia) for 99 mass % TEG and 1 mass % water.

    Example 1:

    Determine the reboiler temperature to regenerate a lean TEG concentration of 99.4 mass % at 101.3 kPa (14.7 psia). If the required reboiler temperature is above the maximum allowable of 204°C (400°F), determine the required partial pressure of a stripping gas (methane) to keep the reboiler temperature at 199°C (390°F).

    Solution:

    Figure 4 indicates that for a lean TEG concentration of 99.4 mass %, the required reboiler temperature without any stripping gas is 220.5°C (428.8°F). Obiviously, the is above the maximum allowable operating reboiler temperature. From the same figure, for a reboiler temperature of 199°C (390°F), the required mole % of striping gas in the vapor phase is 35.6. Therefore, the required partial of stripping gas is:

    Example 2:

    It is desired to regenerate a lean TEG solution to a concentration of 99.4 mass % at 101.3 kPa and 199°C (14.7 psia & 390°F). How much stripping gas is required in a regenerator with 1 theoretical tray?

    Soultion:

    Figure 3A indicates that for a lean TEG concentration of 99.4 mass%, the required striping gas rate is 10.9 Std m3/m3 of TEG (1.44 SCF/gal TEG). The coressponding water and methane partial pressures from Figure 2A are 57.8 and 36.5 kPa ( 8.4 and 5.3 psia), respectively. Note that without any stripping gas, the maximum achievable lean TEG concentration is 99 mass % and the coressponding water and methane partial pressures are 94.6 and 0 kPa (13.7 and 0 psia), respectively. The stripping gas reduced the water partial presuure by 38.9 %.

    Conclusions:

     To achieve a higher lean TEG concentration at a specified reboiler temperature, i.e. below maximum of 204°C (400°F), and pressure, a stripping gas is commonly used. In this TOTM, the mechanism of stripping gas was reviewed. The impact of stripping gas on lean TEG solution concentration or boiling (bubble point) temperature was further investigated quantitatively. The impact of stripping gas on regeneration of TEG solution concentration can be summarized as:

    • Each constituent of a mixture exerts its own partial pressure as a function of temperature and composition.
    • Total system pressure is the sum of all partial pressures.
    • For a fixed pressure, stripping gas lowers the partial pressure of water in the vapor phase
    • Even though this system is not ideal, from Raoult’s Law (equation 1) at a fixed pressure and temperature the concentration of water in the TEG solution decreases as the partial pressure in the vapor phase decreases

    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

    Reference:

    1. Campbell, J. M., “Gas Conditioning and Processing”, Vol. 2, The Equipment Module, 8th Ed., Second Printing, J. M. Campbell and Company, Norman, Oklahoma, 2002.
    2. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2013/05/teg-dehydration-stripping-gas-charts-for-lean-teg-regeneration/, June 2013.
    3. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2013/06/teg-dehydration-stripping-gas-correlations-for-lean-teg-regeneration/, July 2013.
    4. Smith, J.M., Van Ness, H.C. and M.M. Abbott, “Introduction to Chemical Engineering Thermodynamics,” 7th, McGraw-Hill, New York, 2004.
    5. http://msdssearch.dow.com/PublishedLiteratureDOWCOM/dh_004d/0901b8038004d042.pdf, Dow Chemical Company, 2007.
    6. ProMax 3.2, Bryan Research and Engineering, Inc., Bryan, Texas, 2013.
    7. Soave, G., Chem. Eng. Sci. Vol. 27, No. 6, p. 1197, 1972. 

    Appendix A: Vapor-Liquid Equilibrium

    The criteria for vapor-liquid equilibrium are the equality of fugacity of each component in the mixture between phases. Applying these criteria to fugacity (f), of component (i) in the vapor (V) and liquid, (L), phases:

    For a non-ideal mixture of water, TEG and methane, one can write the following expressions in terms of fugacity coefficients (φ) in the vapor phase and activity coefficients (γ) in the liquid phases.

    And,

    Where:

    Equating equations A2 and A3 and solving for xi:

    Calculation of  involves the use of an equation of state. It is a tedious trial and error procedure; therefore, a computer program should be used. These two terms represent the non-ideality of the vapor phase and if the system is an ideal gas both are set equal to 1. On the other hand, activity coefficient  represents the non-ideality of the liquid phase and is calculated by an appropriate activity coefficient model. For an ideal liquid solution, the activity coefficient is also set to 1.

    Assuming the vapor phase is an ideal gas and the liquid phase is an ideal solution, then equation A4 reduces to equation A5, which is Raoult’s Law.