FACILITIESTips of the Month

2019

Table of Contents

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10

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Optimizing Performance of a Refrigeration System with an

External Sub-Cool Economizer

Impact of Heavy End on the Performance of a Mechanical Refrigeration Plant with MEG

Injection

Design and Operation of Unconventional Surface Facilities:

Process Safety Tips

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60

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Table of Contents

Impact of Liquid Carryover on the Performance of a Mechanical

Refrigeration Plant with MEG Injection

A Short Cut Method for

Evaluating Molecular Sieve Performance

Impact of Temperature Approach of the Heat Exchangers on the

CAPEX and OPEX of a Mechanical Refrigeration Plant with MEG Injection

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OPTIMIZING PERFORMANCE OF A REFRIGERATION SYSTEM WITH AN

EXTERNAL SUB-COOL ECONOMIZER

Continuing the January 2008 [1], May 2008 [2], May 2014 [3], and December 2017 [4] Tips of The Month (TOTM), this tip demonstrates the application of an external sub-cooler to optimize the performance of a mechanical refrigeration system. Specifically, by utilizing a cold process stream we will minimize the compressor power and condenser duty. The details of three typical refrigeration systems are given in Chapter 15 of Gas Conditioning and Processing, Volume 2 [5]. They are referred to as follows:1. A simple refrigeration system (Fig 15.1). 2. A refrigeration system employing one flash tank economizer and two stages of compression (Fig 15.7). Also, view the May 2014 [3] and December 2017 [4] TOTM.3. A Simple Refrigeration System with a sub-cool heat exchange economizer (Fig 15.9).

Also, view the May 2008 [2] and May 2014 [3] TOTM.Figure 1 presents the process flow diagrams for a simple system and the system with a sub-cool heat exchange economizer. As part of a hydrocarbon dew point control plant, this tip will evaluate and compare these two refrigeration systems.Let’s consider cooling the process gas to -20°C [-4°F] by removing 2733 kW (9.325 MMBtu/hr) in a propane chiller and rejecting it to the environment by a propane condenser at 37.8°C [100°F]. Pure propane is used as the working fluid in the simulation. In this tip, all simulations were performed with UniSim Design software [6] using the Peng-Robinson equation of state.

Fig 1. Process flow diagrams for a simple refrigeration system and with a sub-cool economizer

Published January 2019Written by: Dr. Mahmood Moshfeghian

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Assuming an approach temperature of5°C [9°F] and a 6.9 kPa (1 psi) pressure drop in the propane chiller, the pressure of saturated propane vapor leaving the chiller is 203.3 kPa (29.5 psia), and at a temperature of -25°C [-13°F]. Assuming no frictional losses in the suction line to the propane compressor, the resulting suction pressure is 203.3 kPa (29.5 psia).The condensing propane pressure at the specified condenser temperature of 37.8 °C (100 °F) is 1303 kPa (189 psi). The condenser frictional losses, plus the frictional losses in the piping from the compressor discharge to the condenser was assumed to be 34.5 kPa (5 psi); therefore, the compressor discharge pressure is 1338 kPa (194 psia). The propane compressor adiabatic efficiency was assumed to be 75%.

External Sub-Cool, Economizer:The process streams 9A and 9B are part of a hydrocarbon dew point control plant and are shown on the top of Figure 2. This stream is the extracted NGL stream from the refrigeration plant, combined with the plant inlet condensate. The stream properties are shown in Table 1. To prepare the liquids to be fed to the deethanizer, the process specification is to raise the temperature of the NGL product stream 9A from -3.9°C (25°F) to 20°C (68°F) in HEX E-102. The resulting duty is 713.6 kW (2.435 MMBtu/hr). This heat will be supplied from a propane refrigerant sub-cool economizer heat exchanger. The process duty and the temperature of the NGL product stream are set by the stabilization process requirements, thus the sub-cool economizer duty is fixed.The sub-cool economizer cools the

condensed propane (refrigerant stream R4) from 37.8°C (100 °F) at 1303 kPa (189 psia) to a cooler temperature at 1269 kPa (184 psia), depending upon the specified propane refrigerant flow rate (stream R5). The pressure drops in HEX E-102 and HEX E-104 are 35 kPa (5 psi) respectively. The heat removed by the sub-cool economizer is fixed by the process duty required to heat the NGL process stream 9A.

Determination of Refrigerant Circulation Rate of Sub-Cool Economizer System:The refrigerant circulation rate has a considerable impact on the compressor power and consequently on the condenser duty. Figure 2 presents the variation of compressor power as a function of the refrigerant circulation rate. This figure indicates that the power is minimum at 995.5 kW (1335 hp) for a circulation rate of 690 kmol/h (1521.5 lbmol/hr), with fixed propane chiller and sub-cool exchanger duties.

Summary:For the same chiller duty, chiller and condenser temperatures, adiabatic compression efficiency, and pressure drops, the results of the sub-cool exchange economizer system are compared with the results of a simple refrigeration system in Table 2. This table indicates that by utilizing an external sub-cool exchange economizer for this case study with optimized propane circulation rate, the compressor power and condenser duty are reduced by 25 % and 25.4%, respectively.

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Fig 2. Cold process stream 9A part of the hydrocarbon dew point control plant is utilized to sub-cool refrigerant stream R5.

Table 1. Process conditions for streams 9A and 9B

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Fig 3. Impact of refrigerant circulation rate on the compressor power

Table 2. Comparison of the key parameters of two refrigeration systems

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To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Process-ing), G5 (Practical Computer Simu-lation Applications in Gas Process-ing) and G6 (Gas Treating and Sulfur Recovery) courses.

References:1. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2008/01/refrigeration-with-flash-economizer-vs-simple-refrigeration-system/, John M. Campbell Tip of the Month, January 2008.

2. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2008/05/flash-tank-vs-hex-economizer-refrigeration-system/, John M. Campbell Tip of the Month, May 2008.

3. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2014/05/refrigeration-with-heat-exchanger-economizer-vs-simple-refrigeration-system/, PetroSkills Tip of the Month, May 2014.

4. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2017/12/optimizing-performance-of-refrigeration-system-with-flash-tank-economizer/, PetroSkills Tip of the Month, December 2017.

5. Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.

6. UniSim Design R443, Build 19153, Honeywell International Inc., 2017.

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IMPACT OF HEAVY END ON THE PERFORMANCE OF A MECHANICAL REFRIGERATION PLANT WITH MEG

INJECTION

Published February 2019

Written by: Dr. Mahmood Moshfeghian

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Continuing the January 2019 [1] Tip of The Month (TOTM), this tip investigates the impact the heavy end characterizations on the performance of a mechanical refrigeration plant with mono-ethylene glycol (EG or MEG) injection for hydrocarbon dew point (HCDP) control. Specifically, the impact of heavy end characterization on the gas-gas heat exchanger and chiller duties, the mechanical refrigeration system, and the liquid propane recovery will be investigated and reported. The details of a mechanical refrigeration plant with MEG injection and regeneration system are given in Chapters 6 and 15 of the Gas Conditioning and Processing, Volumes 1 and 2 [2, 3], respectively.

Figure 1 presents the process flow diagrams for a typical mechanical refrigeration plant with MEG injection system. In this tip, all simulations were performed with UniSim Design R443 software [4] using the Peng-Robinson equation of state.

Case Study:Let’s consider a rich gas with the compositions and conditions presented in Table 1. Based on the reported molecular weight and relative density for the C7+ fraction, Table 2 presents the estimated normal boiling point (NBP), critical properties and acentric factor which are needed by the equation of state.

Fig 1. Process flow diagrams for a mechanical refrigeration plant using a sub-cool economizer and MEG Injection system

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The objective is to meet a hydrocarbon dew point specification of -20 °C [-4°F] at about 4000 kPa (580 psia) for the sales gas by removing heat in the “Gas/Gas” heat exchanger (HX) with a hot end approach temperature of 5°C [9°F] and in a propane chiller and rejecting it to the environment by a propane condenser (“E-103”) at 37.8°C [100°F]. Pure propane is used as the working fluid in the simulation. The pressure drops in the “Gas/Gas” HX and the propane chiller are assumed to be 34.5 kPa (5 psi).The feed gas is flashed in the “Inlet Separator” at 30 °C (86 °F) and 4000 kPa (580 psia) to remove any condensate. The “Inlet Separator” vapor (stream “2”) is saturated with water by the “Saturate -100” to form stream “2 Wet” upstream of mixing with MEG hydrate inhibitor, stream “EG1” and the recycle stream “18A” from the deethanizer overhead vapor (located at the right hand side of Fig. 1). The estimated hydrate formation temperature of streams “2 Wet” is 14.7 °C (58.4 °F). The hydrate inhibitor is injected at the inlet of “Gas/Gas” HX by stream “EG1” and at the inlet of the “Chiller” by stream “EG2”. Stream “5” cools to about -8 °C (17.6 °F) and stream “7” cools down to the specified temperature of -20 °C (-4 °F) which are below the hydrate formation temperature (HFT) of 14.7 °C (58.4 °F). The injection rates of streams “EG1” and “EG2” for 80 weight % lean MEG and water solution are estimated by the Adjust tool of UniSim. A design margin of 1.1 °C (2 °F) HFT below the cold temperature for streams “5” and “7” were assumed.

Table 3 presents the estimated hydrate inhibition injection rates.Assuming an approach temperature of 5°C (9°F) and a 6.9 kPa (1 psi) pressure drop in the propane chiller (“RefChiller”) shell side, the pressure of saturated propane vapor leaving the chiller is 203.3 kPa (29.5 psia), and at a temperature of -25°C (-13°F). Assuming no frictional losses in the suction line to the propane compressor “K-101”, the resulting suction pressure is 203.3 kPa (29.5 psia).The condensing propane pressure at the specified condenser temperature of 37.8 °C (100 °F) is 1303 kPa (189 psi). The condenser “E-103” frictional losses, plus the frictional losses in the piping from the compressor discharge to the condenser was assumed to be 34.5 kPa (5 psi); therefore, the compressor discharge pressure is 1338 kPa (194 psia). The propane compressor adiabatic efficiency was assumed to be 75%.

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Table 2. Estimated C7+ properties [4]

Table 1. Rich feed gas compositions and conditions

Table 3. Estimated 80 weight % lean MEG hydrate inhibition injection rates

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External Sub-Cool Economizer:The cold Stream 7 is flashed in the 3-phase separator “V-102” at -20 °C (-°4F) and 3931 kPa (570 psia). The vapor stream “4” from this cold separator is used to cool down the incoming warm feed gas in the “Gas/Gas” HX. The heavy liquid stream “8B” (rich MEG solution) from the cold separator is regenerated in the regeneration unit (not shown in Fig. 1) and the lean 80 weight % MEG is recycled and used in streams “EG1” and “EG2”. The cold NGL stream “8” (light liquid phase) from the cold separator, “V-102”, is combined with the plant “Inlet Separator” condensate (stream “3”) in the mixer “Mix-101” to form stream “9” at about 5 °C (41 °F) and 3945 kPa (572.2 psia). To prepare the liquid to be fed to the deethanizer, the process specification is to raise the temperature of the NGL product stream “9A” from about -4°C (25°F) and 1535 kPa (222.6 psia) to 20 °C (68 °F) and 1500 kPa (217.6 psia) in “E-102” HX. The required heat duty will be supplied from a propane refrigerant sub-cool economizer “E-104” HX. The process duty and the temperature of the NGL product stream is set by the deethanizer process requirements, thus the sub-cool economizer duty is fixed.The sub-cool economizer cools the condensed propane (refrigerant stream “R4”) from 37.8°C (100 °F) at 1303 kPa (189 psia) to a cooler temperature at 1269 kPa (184 psia), depending upon the specified propane refrigerant flow rate (stream “R5”). The pressure drops in “E-102” and “E-104” HXs are 35 kPa (5 psi); respectively. The heat removed by the sub-cool economizer is fixed by the process duty required to heat the NGL process stream “9A”.

Deethanizer Specifications and Performance:The deethanizer column specifications are:

►To recover 90 mole percent of propane of the feed in the bottom product and

►Ethane to propane mole ratio equal to 5 % in the bottoms product

►Top and bottom pressures are 1450 and 1500 kPa (210.3 and 217.6 psia); respectively

►Number of theoretical stages 12 plus the condenser and reboiler (determined by the material balance and column shortcut calculations)

The deethanizer simulation results are summarized in Table 4.

Impact of Heavy End Characterization:Figure 2 presents the phase envelopes for the key streams of feed (“Dry Feed”), inlet separator vapor (stream “2”) and sales gas (stream “4”). All phase envelopes are generated on the dry basis. As expected the bubble point curves are very close to each other but large deviations are observed for the dewpoint curves. Similar diagrams for the nC7 and nC8 as the heavy end are presented in the Appendix in Figures 1A and 2A; respectively.

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Table 4. Summary of deethanizer key design parameters for C7+

Fig 2. Phase diagrams for the key streams for the case of C7+ as the heavy end 15

Figures 3, 4, and 5 present the impact of heavy ends on the phase envelope of the key streams of feed, inlet separator vapor (stream “2”) and the sales gas (stream “4”), respectively. These figures indicate that as the heavy components are removed in the “Inlet Separator” and cold separator (“V-102”) from the process streams, the impact of heavy end characterization on the phase envelope reduces and vanishes almost completely for sales gas (stream “4”) in Figure 5.

Fig 3. The impact of heavy end on the phase envelope of the feed stream

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Fig 4. The impact of heavy end on the phase envelope of the inlet separator vapor stream

Fig 5. The impact of heavy end on the phase envelope of the sales gas (Stream 4)

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Table 5 presents the impact of heavy end characterization on the “Gas/Gas” HX and “Chiller” duties. Note that the “Gas/Gas” HX duty is controlled by stream “4” composition and rate. Based on the phase envelopes in Figure 5, the sales gas composition is almost independent of heavy ends because they are removed from the sales gas but the heavy ends have more impact on the composition of streams “2”.Table 5 indicates that as the heavy ends become heavier,►stream “2” flow rate decreases because there is more liquid leaving the “Inlet Separator”.

►stream “4” rate increases by about 0.27% (nC7 to C7+) because most of the C7+ has been removed.

►“Gas/Gas” HX duty is set by stream “4” rate and fixed ΔT=25-(-20) =45 °C (81 °F) because Q = mΔ(HSalesgas – H4).

►“Gas/Gas” HX duty increases slightly, less than 0.8 %, because stream “4” rate increases by about 0.27%

►stream “2A” rate decreases, “Gas/Gas” HX duty increases, stream “5A” gets colder, chiller ΔT decreases; therefore, “Chiller” duty decreases.

Assume the design-heavy end was nC8 and feed gas heavy end is C7+, not the design nC8. More liquids would leave the “Gas/Gas” HX so the chiller duty would decrease by about 38%. But the additional duty to

Table 5. Impact of heavy end on the Gas/Gas HX and Chiller duties

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condense the liquids in the “Gas/Gas” HX has to come from somewhere. If the “Gas/Gas” HX has excess area to accommodate the additional duty requirements, then there would indeed be a decrease in chiller duty. If it does not, the duty of the chiller may actually increase. If the feed gas got lighter, and the heavy end is nC7, not the design nC8, then more gas would go to the chiller (less liquids leaving the “Gas/Gas” HX) and the chiller duty would increase by about 20%. Here, the chiller would have to have excess capacity.

This indicates that a change in feed gas characterization would have an effect on the ability of a refrigeration unit to make spec. For easier reference of the stream, see Figure 6.

Fig 6. Simplified schematic of the front end segment of the process flow diagram

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Table 6. Impact of heavy end on the refrigeration systems key parameters

Table 6 presents the impact of heavy end characterization on the refrigeration systems. This table indicates that the rate, compressor power, condenser and the sub-cool economizer duties decrease as the heavy end becomes heavier. Table 6 also indicates that the rate, compressor power and condenser duty for the sub-cool economizer refrigeration system are lower compared to the simple refrigeration system. Because the chiller duty decreases, the refrigeration systems become smaller; therefore, the OPEX and CPEX decrease.The heat removed by the sub-cool economizer “E-104” is used to heat stream “9A” in “E-102” HX. Location of “E-102” HX and streams “9A” and “9B” are shown in Figure 7.

Table 7 presents the impact of heavy ends on the rates and the molecular weights for stream “3” from the “Inlet Separator” and stream “8” from the cold separator (“V-102”) and the combined NGL stream “9”. This table indicates that as the heavy end becomes heavier, rate of stream “3” increases but the rate of stream “8” decreases. Because the rate of heavy ends entering the “Gas-Gas” HX and “Chiller” decrease, the chiller duty decreases and condensation of components decrease resulting lower streams “8” and “9” rates. Table 1A in the Appendix present components flow rates for streams “3” and “8”.

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Table 6. Impact of heavy end on the refrigeration systems key parameters

Table 7. Impact of the heavy end on streams “3” and “8” and the combined NGL stream “9” rates and molecular weight

Fig 7. Simplified schematic of the back end of the process flow diagram

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Using “VLV-100” stream 9 pressure is reduced from 3945 kPa (572.2 psia) to 1535 kPa (222.6 psia) in stream “9A”. Table 8 presents the combined NGL streams “9A” and “9B” (see Figure 7) properties. This table indicates that as heavy end becomes heavier, the “E-102” HX duty decreases because combined NGL stream rate decreases. The required heat for this HX is supplied by the sub-cool HX (“E-104”) of the refrigeration system.

Table 8. Impact of the heavy end on the combined NGL streams “9A” and “9B” properties

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Table 9 presents the impact of heavy end on the plant overall material balance. This table indicates that as the heavy end becomes heavier,

►the sales gas rate increases (stream 4)

►the deethanizer feed (combined NGL stream, 9 ) rate decreases because the sales gas rate (stream 6) has increased

►the overhead vapor temperature from the deethanizer top remains almost constant because the overhead composition does not significantly change

►the overhead vapor rate from the deethanizer top decreases because the deethanizer feed rate decreased.

Table 9. Impact of the heavy end on liquid propane recovery

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The overhead vapor of deethanizer is compressed from 1450 kPa (210.3 psia) to the feed gas inlet pressure of 4000 kPa (580 psia) by the recycle compressor (“K-100”) and cooled down to the inlet feed gas temperature of 30 °C (86 °F) in the “E-101” HX. The liquid from compressor suction scrubber is recycled and combined with deethanizer feed by the recycle pump. Table 10 presents the compressor and pump power and the “E-101” HX duty requirements. Table 10 indicates that as the heavy end becomes heavier, the recycle compressor and pump power and the cooler duty decrease because the recycle stream rates decrease.

Summary:The feed analysis and /or heavy end characterization in natural gas play an important role in the equipment sizing and process design. Feed analysis may change when different wells of slightly different composition are brought to the production facility. This tip demonstrated the impact of heavy end characterization in the feed gas on the process streams rates, phase behavior, the equipment sizes and the refrigeration requirement by replacing, the C7+ with n-heptane (nC7) and n-octane (nC8). All other specifications and operating conditions were kept the same.

As demonstrated in this tip, it would be a good practice to size the equipment with a design margin of 1.2 to 1.3 to take into account the changes in feed gas heavy end composition and characterizations.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing) and G6 (Gas Treating and Sulfur Recovery) courses.

Table 10. Impact of the heavy end on the recycle compressor, pump, and cooler

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References1. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2019/01/optimizing-performance-of-refrigeration-system-with-an-external-sub-cool-economizer/, PetroSkills -John M. Campbell Tip of the Month, January 2019.

2. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum

3. Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.

4. UniSim Design R443, Build 19153, Honeywell International Inc., 2017.

Continuing the August 2018 Tip of the Month (TOTM) [1] on design and operation of unconventional surface facilities, this TOTM presents process safety tips for four case studies:

1. Direct Fired Heater Treater Burn Through Failures2. Tank Blanket Gas / Flame Arrestors3. Pocketing Vent / Relief Piping4. Hot Oiling of Oil Storage Tanks to meet TVP / RVP We start this tip with a quote from a colleague, James A. Britch: “I never regretted buying quality”. There is a lesson in there for unconventional batteries. There is tremendous pressure to reduce capital costs, but you should be focused on life-cycle cost. If you install equipment and then burn down the battery…you haven’t saved much. There’s the loss of capital and revenue from shut-in production.

Published March 2019

Written by: James F. Langer, P.E.

DESIGN AND OPERATION OF UNCONVENTIONAL SURFACE

FACILITIES: PROCESS SAFETY TIPS

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This Tip of the Month is also available as an on-demand webinar.

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Process Safety Case Study 1: Direct Fired Heater Treater Burn Through FailuresDirect fired burners are failing due to internal flame impingement directly on the steel, and salt build up on the outside of the firetube in the process fluids. The salts build up, act as an insulator, and then steel temperature increases until a burn through occurs. Since the process side of the burner operates at a higher pressure than the natural draft burner, the process fluids enter the burner and ignite. In many instances, this has resulted in massive damage to the entire battery.Figure 1 illustrates the location of burn through failure that occurs in the 12 o’clock portion of the burner firetube (combustion leg).Solutions to this issue are to use:(1) Indirect fired heater where the process fluid flows through coils in a larger shell surrounded by a heat transfer fluid. The firetube burner is immersed in the heat transfer fluid within the same shell.

a. This solves the salt build up failure mode, but not the flame impingement failure. b. Some operators use glycol as a heat transfer fluid. This is a cheaper alternative to expensive heat transfer fluids, but glycols degrade to acetic acid if the skin temperatures of the fire tubes exceed 350 °F (177 °C). These same operators rarely check the pH of their heat transfer fluid until leaks and severe corrosion are found. Heat transfer oils are a better choice.

(2) Direct fired heater with a ceramic sleeve to take the higher temperature shock of direct flame impingement. Heaters normally have ceramic refractory castables and bricks to prevent direct flame impingement. As shown in Figures 2 and 3 Bartz et al. [2] recommends inserting a ceramic tube to spread the heat flux of the flame and avoid burn through.

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Fig 2. ALZETA distributed flux burner [2]

Fig 1. Heater Treater Firetube Failure

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(3) Best Solution: consider a separate furnace and heat exchanger using heat transfer oils. This solves both issues. The heat transfer oils are designed to operate without degradation at process temperatures required, and the metal temperatures in the heat exchanger will not cause metal failure if there are no salt deposits on the heat exchanger. This may require periodic hydro-blasting, but you will not have to rebuild the battery.

Fig 3. ALZETA general arrangement of distributed flux burner and blower [2]

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(4) Think Reliability has a free excellent Root Cause Analysis Excel tool (Figure 4).

Fig 4. CAUSE MAP for Burner Firetube Failures in Heater Treaters

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Process Safety Case Study 2: Tank Blanket Gas / Flame Arrestors [3]Flame arrestors and tank blanket gas provide independent layers of protection between the ignition source - flare or thermal oxidizer - and the vapor space in the water and oil tanks. When evaluating what to use in a design consider using a layer of protection analysis or LOPA [3]. This tool is discussed in the PetroSkills-John M Campbell PS4-Process Safety Course. It provides a semi-quantitative solution to design decisions that are based of failure frequencies and not just personal preferences or gut feels. The more independent layers of protection the lower the frequency of the consequence occurring.

f = (IEF) x PFD1 x PFD2….

Where:f = Frequency of the consequence occurring for the scenario

IEF = Frequency of the initiating event

PFD = Probability of the failure on demand for an independent layer of protection.For example, the probability that a relief valve will not operate as intended.

Flame arrestors have a Probability of Failure on Demand (PFD) with a range of 1x10-1 for arrestors without

temperature indicators and an effective isolation / shutdown system, and tank blanket gas (BPCS-Basic Process Control System) has a PFD of 1x10-1 [3]. The designer can use both or either to provide independent layers of safety protection. Flame arrestors are subject to plugging from ice, corrosion, fouling, improper or lack of maintenance. Blanket gas works well in this situation due to the narrow range of flammability of methane in the air (5-15% fuel to air). The majority of stock tank incidents occur during maintenance activities with small amounts of gas and large volumes of air. During the high volume of production timeframe for unconventional tank batteries the stock tanks degas and have “auto-blanketing” of the active tank vapor space. But what happens in the future when rates are very low? What happens in water tanks that are not provided with gas blanketing? This also explains why water production tanks / injection batteries experience tank fires/explosions. In general, methane has a very low solubility in water - approximately 2 SCF/STB (0.36 Sm3/STm3) of water going from 250 psi (1724 kPa) to atmospheric pressure. This small volume often results in flammable mixtures in the vapor space of the tank.It is a good practice to select tank blanketing as your first line of defense to prevent internal tank corrosion and gas plant amines and TEG process corrosion/solution degradation by keeping oxygen out of the system, and internal explosions from flash back from flares and thermal oxidizers.

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Process Safety Case Study 3: Pocketing Vent / Relief PipingMany unconventional tank batteries run their vent / flare / thermal oxidizer piping at grade on sleepers with zero slope, then jump vertically into a flare knockout (see Figure 5). This pocketed piping is a liquid trap for water and heavier hydrocarbons. Once a liquid pocket forms, the tanks overpressure, and then vent locally through their pressure/vacuum reliefs and thief hatches. In the winter this pocket can freeze and block the flare (see Figure 6). This can also cause a loss of containment when a PSV activates and cannot depressure to the flare. This causes a loss of revenue, as well as an environmental and safety issue. These vapors are extremely rich and normally are much heavier than air. This creates the potential for an unconfined vapor cloud explosion or flash fire locally.

Fig 6. This flare knock out has a 5 ft (1.5 m) pocket - tanks vent, and the potential for ice blockage in the winter

Fig 5. Pocketed piping causes an overpressure in the tanks and results in venting

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As illustrated in Figure 7, slope your vent/relief piping toward a lower elevation knock out or burn pit.API Standard 521 [4], pressure-relieving and depressuring systems, requires that the flare piping be free draining to the flare knock out drum, and then free draining from the flare back to the flare knock out.

These issues causing tanks to vent heavier than air molecules (propane/butane) can lead to flash fires and unconfined vapor cloud explosions (UCVE). Heavy vapor generally finds an ignition source. Figure 8 shows how oxygen may get into the oil stock tanks. Getting oxygen into your system causes major damage to the gas plant amine systems, and TEG systems, as well as general corrosion in your facilities.

Fig 7. Flare header design into knockout – sloped / no pockets

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Fig 8. How oxygen gets into your oil stock tanks & also causes venting

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Process Safety Case Study 4: Hot Oiling Oil Storage TanksAs illustrated in Figure 9, some operators use hot oil trucks during winter months to heat the crude oil in the tanks to flash light ends off the crude to meet vapor pressure specifications for crude sales. The solution to this issue is not to use hot oil trucks but is to stabilize the crude or use a design using Vapor Recovery Towers (VRT) as discussed in the August 2018 Tip of the Month- Design and Operation of Unconventional Surface Facilities Issues-Stabilization [1].

Many operators are faced with large numbers of tank batteries in the hundreds or thousands spread over a large geographic area. Many rely on unsupervised contractors to conduct hot oiling operations, oil & water loading and unloading, and other maintenance operations.►Do you have an operator present at the location to help with hot oiling?

►Does the lease operator visit the site with the contractor to issue a Hot Work Permit and JSA?

►Do you have operating guidelines or checklists for Hot Oil Operations?

►Do they include monitoring of weather conditions? Wind Speed? - Shutdown for low or no wind conditions?

►Do they include stopping of other operations such as oil / water truck loading?

►Are parts of the lease blocked off to prevent other vehicles entering as ignition sources?

►Are the hot oil trucks placed upwind of the tanks?

Fig 9. Winter “Hot Oiling” of oil stock tanks

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►Are the hot oil trucks 100 ft or 50 ft from the tanks? What’s your company’s design spacing requirements for ignition sources / direct fired equipment and oil storage tanks?

►Have you conducted a HAZOP for Hot Oiling Operations? Many accidents happen during non- routine operations.

►Do you have gas detectors? Normally the contractor has a portable. Is that effective? You have RVP issues because it’s winter and cold. Where will the contractor be…in the truck where it’s warm.

The hot oil truck is a direct fired (propane) heater with propane storage, and diesel or oil storage. It is normally used to pump hot oil at high pressure down the well’s tubing to melt wax deposits. This operation is normally done for 24 hrs/day during winter months to stabilize the crude. It is extremely dangerous, and many flash fires have occurred in the past few years.February 2, 2018 Explosion-Flaming Truck ignites Tank Battery-worker airlifted with serious burns These are just some recent examples, but unfortunately, there are many, many more…

Read: August 24, 2018 Fire Crews Battle Tank Battery Fire

Read: January 25, 2019 Officials identify victims of fire on FM 1788 South

Read: January 2, 2019 Noble Energy oil tanks in West Texas catch fire

Watch Video: Heater Treater FireMany Tank Battery Fires are occurring in areas with new unconventional plays like the Bakken and West Texas.

Summary and ConclusionsIn this tip of the month, we have identified process safety risks with some designs, solutions to the issues, and evidence that the problem exists.So my question for you….after reading this tip, what action do you take to improve the safety of your designs and operations at your company? Your company and colleagues need you to take action.►Are your designs safe?►Are your operations safe?Stay Safe! Let us know if you have any questions.To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), PF3 (Concept Selection and Specification of Production Facilities in Field Development Projects), PF4 (Oil Production and Processing Facilities), PF49 (Troubleshooting Oil & Gas Processing Facilities), and PS4 (Process Safety Engineering) courses.

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References1. Langer, J.F., https://www.petroskills.com/blog/entry/00_totm/aug18-fac-design-and-operation-of-unconventional-surface-facilities, PetroSkills-John M. Campbell Tip of the Month, August 2018.

2. “SPE 166261 Distributed-Flux Burners Improve Life of Firetubes and Process Throughput in Heater Treaters”, David Bartz, Michael Silberstein, James Gotterba; ALZETA Corporation, 2013

3. “Layer of Protection Analysis-Simplified Process Risk Analysis”, Center for Chemical Process Safety-CCPS, 2001, AIChE, Table 5.2 In-line deflagration arrestor

4. API Standard 521, Pressure-relieving and Depressuring Systems. 6th Edition, Jan 2014

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IMPACT OF LIQUID CARRYOVER ON THE PERFORMANCE OF A

MECHANICAL REFRIGERATION PLANT WITH MEG INJECTION

Published April 2019

Written by: Dr. Mahmood Moshfeghian

Problems in meeting sales-gas dew point specifications are not unusual in plants. A facility engineer often suspects separator carryover when troubleshooting such a plant. Proper sizing of equipment for gas-liquid separation is essential to almost all processes. Many facility operating problems are related to improperly designed or under-sized gas-liquid separators. The following list presents items that can contribute to too much liquid (carryover) in the gas stream.►The mist extractor operating Ks value is greater than the design value.

►The velocity profile through the mist extractor is poor, resulting in localized high velocities/flooding.

►The droplet sizes reaching the mist extractor are too small.

►The entrained liquid load reaching the mist extractor is too high.

►The mist extractor is damaged or plugged.

►Level control and instrumentation malfunction or failure

►Foaming

Continuing the December 2005, January and February 2019 [1, 2, 3] Tips of The Month (TOTM), this tip

investigates the impact of the liquid carryover (LCO) on the performance of a mechanical refrigeration plant with mono-ethylene glycol (EG or MEG) injection for hydrocarbon dew point (HCDP) control. Specifically, the impact of LCO on the gas-gas heat exchanger and chiller duties, the mechanical refrigeration system, and the liquid propane recovery will be investigated and reported.The details of a mechanical refrigeration plant with MEG injection and regeneration system are given in Chapters 6, 15 and 18 of the Gas Conditioning and Processing, Volumes 1 and 2 [4, 5], respectively. In addition, how to minimize the liquid carry in separation equipment are discussed in PetroSkills-John M. Campbell course titled, “PF42 - Separation Equipment - Sizing and Selection”. Figure 1 presents the process flow diagrams for a typical HCDP control plant using mechanical refrigeration with MEG injection system. This figure is similar to the February 2019 TOTM [3] with the exception that the refrigeration system utilizes a flash tank economizer with two stages of compression. In this tip, all simulations were performed with UniSim Design R443 software [6] using the Peng-Robinson equation of state.

Case Study:Let’s consider the same case presented in February 2019 TOTM [3] for a rich gas with the compositions and conditions presented in Table 1 [3].

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Fig 1. Process flow diagrams for a HCDP plant using mechanical refrigeration with a flash tank economizer and MEG Injection system

Based on the reported molecular weight and relative density for the C7+ fraction, Table 2 presents the estimated normal boiling point (NBP), critical properties and acentric factor which are needed by the equation of state. The objective is to meet a hydrocarbon dew point specification of -20 °C (-4 °F) at about 4000 kPa (580 psia) for the sales gas by removing heat in the “Gas/Gas” heat exchanger (HX) with a hot end approach temperature of 5°C (9°F) and in a propane chiller, 5 °C (-4 °F) approach temperature, and rejecting it to the environment by a propane condenser (AC-100) at 37.8°C (100°F). Pure propane is used as the working fluid in the simulation. The pressure drops in the “Gas/Gas” HX and the propane chiller are assumed to be 34.5 kPa (5 psi).

Table 1. Rich feed gas compositions and conditions

Table 2. Estimated C7+ properties [4]

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The feed gas is flashed in the “Inlet Separator” at 30 °C (86 °F) and 4000 kPa (580 psia) to remove any condensate. The “Inlet Separator” vapor (stream 2) is saturated with water by the “Saturate -100” to form stream “2 Wet” upstream of mixing with MEG hydrate inhibitor, stream “EG1” and the recycle stream “18A” from the deethanizer overhead vapor (located at the right-hand side of Fig. 1).The estimated hydrate formation temperature of streams “2 Wet” is 14.7 °C (58.4 °F). The hydrate inhibitor is injected at the inlet of “Gas/Gas” HX by stream “EG1” and at the inlet of the “Chiller” by stream “EG2”. Stream “5” cools to about -8 °C (17.6 °F) and stream “7” cools down to the specified temperature of -20 °C (-4 °F) which are below the hydrate formation temperature (HFT) of 14.7 °C (58.4 °F). The injection rates of streams “EG1” and “EG2” for 80 weight % lean MEG and water solution are estimated by the Adjust tool of UniSim. A design margin of 1 °C (1.8 °F) HFT below the cold temperature for streams “5” and “7” were assumed.Assuming an approach temperature of 5°C (9°F) and a 6.9 kPa (1 psi) pressure drop in the propane chiller (“Chiller”) shell side, the pressure of saturated propane vapor leaving the chiller is 203.3 kPa (29.5 psia), and at a temperature of -25°C (-13°F). Assuming no frictional losses in the suction line to the propane compressors “K-101” and “K-102”, the resulting suction pressure is 203.3 kPa (29.5 psia).The condensing propane pressure at

the specified condenser temperature of 37.8 °C (100 °F) is 1303 kPa (189 psi). The condenser “AC-100” frictional losses, plus the frictional losses in the piping from the compressor discharge to the condenser were assumed to be 34.5 kPa (5 psi); therefore, the discharge pressure of compressor “K-102” is 1338 kPa (194 psia). The compressors inter stage pressure was determined by equalizing the power for “K-101” and “K-102”. The compressors adiabatic efficiency was assumed to be 75%.The cold Stream 7 is flashed in the 3-phase separator “V-102” at -20 °C (-°4F) and 3931 kPa (570 psia). The vapor stream “4” from this cold separator is used to cool down the incoming warm feed gas in the “Gas/Gas” HX. The heavy liquid stream “8B” (rich MEG solution) from the cold separator is regenerated in the regeneration unit (not shown in Fig. 1) and the lean 80 weight % MEG is recycled and used in streams “EG1” and “EG2”. The cold NGL stream “8” (light liquid phase) from the cold separator, “V-102”, is combined with the plant “Inlet Separator” condensate (stream “3”) in the mixer “Mix-101” to form stream “9” at about 5 °C (41 °F) and 3945 kPa (572.2 psia). To prepare the liquid to be fed to the deethanizer, the process specification is to raise the temperature of the NGL product stream “9A” from about -4°C (25°F) and 1535 kPa (222.6 psia) to 20 °C (68 °F) and 1500 kPa (217.6 psia) in “E-102” HX. The process duty and the temperature of the NGL product stream is set by the deethanizer process requirements. The pressure drops in “E-102” HX is 35 kPa (5 psi).

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Deethanizer Specifications and Performance:Like the February 2019 TOTM [3], the deethanizer column specifications are: A. To recover 90 mole percent of propane of the feed in the bottom product and B. Ethane to propane mole ratio equal to 5 % in the bottoms product C. Top and bottom pressures are 1450 and 1500 kPa (210.3 and 217.6 psia); respectively D. Number of theoretical stages 12 plus the condenser and reboiler (determined by the material balance and column shortcut calculations)The deethanizer simulation results are summarized in Table 3.

Table 3.

Summary of deethanizer key design parameters

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Impact of Liquid Hydrocarbon Carryover:Separator “V-102” is a three-phase separator. Under ideal condition the vapor (stream 4) leaving the separator has no LCO and its dewpoint temperature is the same as the feed (stream 7) temperature. Typical range of liquid carry over is 0.013–0.27 m3 liquid/106 std m3 of gas (0.1–2 gallon of liquid/MMscf) [5]. In practice due to the reasons listed in the preceding section the LCO can be even higher. In this tip, the impact of LCO was investigated for a range of 0 to 3 mole % of liquid in light liquid phase (liquid hydrocarbon phase) entrained into the gas phase. The entrained liquid consists of heavier molecules causing the dewpoint temperature of

stream 4 and sales gas to go up and make it off spec. To offset the effect of LCO and bring back the sales gas dewpoint temperature to spec, the operators typically lower the chilling temperature of feed (stream 7) to separator (“V-102”). This is possible if the mechanical refrigeration system is capable of handling a higher chilling load.Figure 2 presents the hydrocarbon dewpoint curves as a function of the liquid hydrocarbon carryover (CO). the cricondentherm points shift to the right as LCO increases. The bubble point curves are not presented because the LCO has negligible effect on the bubble curves. All phase envelopes are generated on the dry basis

Fig 2. Impact liquid carryover on the sales gas hydrocarbon dewpoint temperature

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Figure 3 presents the impact of LCO on the sales gas dewpoint temperature and the required cold separator feed (chilling) temperature to offset the LCO. As the LCO increases the chiller temperature should be decreased to meet the sales gas dewpoint spec of -20 °C (-4 °F). For 3 mole % LCO, the sales gas dewpoint temperature is -14.4 °C (6.1 °F). To bring back the sales gas dewpoint temperature, the process gas (stream 7) should be cooled to -28.6 °C (-19.5 °F).

Fig 3. Impact of liquid carryover on the sales gas dewpoint temperature (solid line) and the required cold separator feed temperature (dashed line) to offset liquid carryover

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As the chiller temperature decreases, to counter the effect of the LCO, the hydrate formation temperature depression of streams 5 and 7 increases, which requires a higher MEG injection rate. Figure 4 presents the impact of LCO on the rate of streams EG1 and EG2 upstream of Gas/Gas HX and the chiller, respectively. Note the required inhibitor injection rate for stream EG2 upstream of the chiller increases considerably with the increase in LCO.If there is carryover of the hydrocarbon phase, there is also

a likelihood of carryover of the glycol phase. This can result in problems meeting the water dewpoint specification and also introduces a deleterious substance into the sales gas (MEG) which may also not be allowed in the sales gas contract. Lowering the chiller temperature to counter the effect of LCO also cause an increase in the compressor power, Gas/Gas-gas HX, chiller and condenser duties. Figures 5 A and B illustrate the impact of LCO on the compressor power and Gas-Gas HX, Chiller, and condenser duty in SI and FPS system of units, respectively.

Fig 4. Impact of liquid carryover on the MEG injection rate upstream of Gas/Gas HX (EG1) and chiller (EG2)

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Fig 5A. Impact of liquid carryover on the compressor power, Gas-Gas HX, chiller, and condenser duty

Fig 5B. Impact of liquid carryover on the compressor power, Gas/Gas HX, chiller, and condenser duty

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Figure 6 presents the impact of LCO on the liquid propane and sales gas recoveries. This figure indicates that as the LCO increases from 0 to 3 mole %, the liquid propane recoveries increase from about 17% to 27% on mole basis but the sales gas recovery decreases slightly from about 97 % to 96 % on mole basis. The extra propane liquid recovery is achieved by operating the chiller at a lower temperature which requires higher OPEX and CAPEX. As the chiller temperature is reduced, more ethane and methane end up in the low temperature separator (LTS), V-102, liquids. These need to get boiled out in the deethanizer, so the duty of E-102 increases, reboiler and condenser duty of the deethanizer increase, and the recompression power in K-100 also increases. It may also be possible to flood the deethanizer.The cold condenser on the deethanizer requires propane for cooling. These units are also designed with no condenser. The cold LTS (V-102) liquid is used as reflux, and the

liquids from the inlet separator are introduced lower in the column. V-100 is not required then. E-102 is usually a feed/bottoms heat exchanger. Overhead of the deethanizer can likely go to the sales gas and may not have to be recycled. It really should not contain anything heavier than propane, heavy key (HK).Table 4 presents the impact of LCO on the key equipment incremental capacity requirement to meet the sales gas hydrocarbon dewpoint temperature by lowering the chiller temperature. Assume the system was built with a design margin factor of 1.25. Table 4 indicates that this system can handle up to one mole % LCO with higher OPEX. However, for more than one mole % LCO, the system cannot lower chiller temperature enough to meet the sales gas dewpoint temperature. As shown in Table 4, the compressor power, hydrate inhibition rate, condenser and chiller duties are the limiting factors. Under such condition it may require plant shutdown for trouble shooting to reduce the LCO.

Table 4. Estimate of equipment incremental capacity requirement to handle liquid carryover

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Summary:The common practice to meet sales gas hydrocarbon dewpoint temperature under the condition of liquid carry is to operate the chiller at a temperature below the sales gas hydrocarbon dewpoint spec. This is only possible if the key equipment can handle the extra load with higher OPEX. This tip demonstrated the impact of varying the LCO from 0 to 3 % on a mole basis on the process stream rates, phase behavior, the equipment sizes and the refrigeration requirement. As demonstrated in this tip, it would be a good practice to size the equipment with a design margin of 1.2 to 1.3 to consider the changes in operation conditions and the liquid carryover. Most important to minimize LCO is to have a properly designed separator with good feed pipe, inlet device, mist extractor, gas gravity separation and liquid gravity separation sections. To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), and PF42 (Separation Equipment - Sizing and Selection) courses.

References:1. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2005/12/impact-of-liquid-carry-over-on-sales-gas-dew-point/, PetroSkills -John M. Campbell Tip of the Month, December 2005.

2. Moshfeghian, M., https://www.petroskills.com/blog/entry/00_totm/jan19-fac-optimizing-performance-of-refrigeration-systems-with-an-external-sub-cool-economizer, PetroSkills -John M. Campbell Tip of the Month, January 2019.

3. Moshfeghian, M., https://www.petroskills.com/blog/entry/00_totm/feb19-fac-impact-of-heavy-end-on-the-performance-of-a-mechanical-refrigeration-plant-with-meg-injection, PetroSkills -John M. Campbell Tip of the Month, February 2019.

4. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum

5. Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.

6. UniSim Design R443, Build 19153, Honeywell International Inc., 2017.

A SHORT CUT METHOD FOR EVALUATING MOLECULAR SIEVE

PERFORMANCE

Published May 2019

Written by: Harvey H. Malino, P.E.

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This Tip of the Month shows how a Short Cut Method (SCM), after one Performance Test Run (PTR), may be used to estimate the life of a Type 4A molecular sieve dehydrating a water-saturated feed of natural gas. The May 2015 Tip of the Month [1] discussed the benefits of standby time. In that Tip (which the reader is urged to revisit), a case study was presented for a 3-tower dehydration system. The system was designed to meet a three-year life; however, a PTR after one year of service predicted the total life of the molecular sieve would only be about two years. Using the available stand-by time, the molecular sieve life was extended to about 3.7 years.

The above results were calculated used the concepts outlined in Chapter 18 of Gas Conditioning and Processing, Volume 2: The Equipment Modules (9th Edition) [2]. Due to the non-linearity of the modeling techniques, the manual calculations are tedious. The June 2016 Tip of the Month [3] showed how a computer module developed for the PetroSkills-John M. Campbell GCAP (Gas Conditioning and Processing Software) [4] could be used to perform the same calculations. Among other things, the GCAP module directly calculates the life factor, FL, which generates more consistent results compared to visually reading the Generic Molecular Sieve Decline Curves (Figure 1).

Fig 1. A generica molecular sieve decline curves [1]51

Short Cut Method (SCM):By making a few simplifying assumptions, manual calculations can be performed on the back of an envelope. This SCM permits an operator or a facility engineer to quickly determine the life expectancy of the molecular sieves. The assumptions, which cover the majority of natural gas dehydration units in the field, include:1. Water-saturated natural gas is the feed

2. The feed conditions remain relatively constant throughout the life of the molecular sieve system

3. Type 4A-1/8 inch (3 mm) pellets or 4x8 beads are used

4. The fresh equilibrium loading is 23 weight percent water

5. The residual loading is 4 weight percent water

6. 5 % of the bed weight is devoted to the Mass Transfer Zone (MTZ)

7. Normal life degradation following the curve shapes in Figure 1. If upsets such as liquid carryover, bed lifting, bed support failure, valve hang-ups, contaminants in the regeneration gas, flow channeling or other adverse conditions occur, the shape of the

curves in Figure 1 will be quite different.Chapter 18 of Gas Conditioning and Processing, Volume 2: The Equipment Modules (9th Edition) [2] contains equations that permit us to calculate the total mass of the molecular sieve, the Break Through Loading (BTL, or Useful Loading), and the aged net equilibrium loading. Using the assumptions listed above together with the equations in Chapter 18 it can be shown that:

FL = BTL/18 (1)where:FL = Life factor

BTL = (100)(mass of water removed/mass of molecular sieve) (2)

The remainder of this Tip of the Month will compare the results of the SCM to those obtained from the rigorous manual method and the computer-generated method.

SCM vs Rigorous Method:Figure 2 shows the Process Flow Diagram used for the Case Study [2]. Tables 1 gives the Design Basis. Table 2 presents the Design Summary. Table 3 shows the Results of the PTR after one year of operation (note the feed flow rate and the temperature during the PTR are slightly different than the design basis).

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Fig 2. Typical process flow diagram for a 3-tower adsorption dehydration system [2]

Table 1. Design basis for the case study

Table 2. Design Summary for the case study

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Additional information used in the Case Study:►3 tower system (2 towers on adsorption, 1 on regeneration)

►External Insulation

►Tower ID = 2.9 m (9.5 ft)

►Each tower contains 24630 kg [54300 lbm] of Type 4A 4x8 mesh beads and is designed to last three years.

►Regeneration circuit capable of handling an extra 15% of flow

►Unit is operated on fixed time cycles

►No step-change events such as liquid carryover, poor flow distribution, etc.

Following is a recipe for using the SCM:1. Use Equation 2 to calculate the design BTL = 10.6 wt %►BTL =100 (16 h)(163 kg water removed/h)/(24630 kg mol sieve) =10.6%►BTL =100 (16 hr)(360 lbm water removed/hr)/(54300 lbm mol sieve) =10.6%

2. Use Equation 1 to calculate the design FL = 10.6/18 = 0.59

3. Locate the design FL at 1095 cycles on Figure 3, Calculated Life Factors (FL=0.59 & 3 years of 24-hour cycles per tower is equivalent to 1095 cycles). This is the Design FL Point.

4. Use Equation 2 to calculate the PTR BTL = 12.0 wt %►BTL =100 (20.9 h)(141 kg water removed/h)/(24630 kg mol sieve) =12%

Table 3. Results of Performance Test Run (PTR) after 12 months of operation

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Table 3. Results of Performance Test Run (PTR) after 12 months of operation►BTL =100 (20.9 hr)(312 lbm water removed/hr)/(54300 lbm mol sieve) =12%

5. Use Equation 1 to calculate the PTR FL = 12/18 = 0.67

6. Locate the PTR FL at 365 cycles on Figure3, Calculated Life Factors (FL=0.67 & one year of 24-hour cycles per tower is equivalent to 365 cycles). This is the PTR FL Point.

7. Because the PTR FL Point falls on a curve lower than the Design FL Point, we need to be concerned. Interpolating and extrapolating the capacity decline curve from the PTR FL Point, we see an FL of 0.67 (the Design FL) will occur after a total of about 750 cycles. This is approximately one year shorter than the expected life.

Fig 3. Calculated Life Factors

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Because the unit has a regeneration circuit that can handle an additional 15% of flow, the complete regeneration cycle (heating, cooling, de- and re- pressurization) can be reduced to 7.0 hours. This allows the beds to turn around faster. Using the reduced cycle time (the complete cycle time is now 21 hours vs the original 24 hours) and the original design basis conditions with the SCM recipe:1. Calculate End of Life BTL = 9.3 wt % (from Equation 2).►BTL =100 (14 h)(163 kg water removed/h)/(24630 kg mol sieve) =9.3%►BTL =100 (14 hr)(360 lbm water removed/hr)/(54300 lbm mol sieve) =9.3%

2. Calculate End of Life FL = 9.3/18 = 0.52 (from Equation 1). This is because less water is being adsorbed per cycle.

3. Interpolating and extrapolating from the PTR FL Point, we find the End of Life FL of 0.52 occurs around 1400 cycles (see Figure 4). Since 365 cycles have already occurred and going forward a reduced cycle time will be used, the molecular sieves are forecast to last a total of 3.5 years.

Table 4 compares the results of the SCM to the rigorous manual calculations (May 2015 Tip of the Month) and the computer-generated

calculations (June 2016 Tip of the Month). The difference between the predicted life using standby time shown by the Computer Model and the two manual methods is primarily due to the inherent inaccuracy of trying to interpolate and extrapolate data plotted on Figures 3 and 4. The computer model will produce the same result every time. This cannot be said when visually reading Figures 3 and 4. Figure 5 shows the output from the GCAP Computer Model [3]. Note that when working on one curve, the higher the calculated EOL FL, the fewer the Number of Cycles (NOC) remaining until the beds need to be replaced.The PTR in the above Case Study was run after 365 cycles. The slope of the curve is fairly steep in this region and small changes in data can have a significant impact on the life predictions. While the user can get a good indication of the state of decline of their molecular sieve unit, scheduling additional PTR’s is highly recommended. Finally, generic curves are used in these Figures. The shape of your specific molecular sieve capacity decline curve may differ from these generic curves. Despite these caveats, the SCM offers the user a quick and easy way to assess the capacity decline of their molecular sieve unit.

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Fig 4. Calculated Life Factors Using Standby Time

Table 4. Comparison of Three Methods

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Fig 5. Projected life factor (LF = 54.3% and NOC = 1251.4) if standby time is used

Summary:We can draw the following conclusions from this case study:

►The short-cut method presented allows the user to quickly estimate the decline of their adsorbent based on only one performance test run for molecular sieve dehydrators using low-pressure regeneration. This permits the early formulation of a credible action plan. The short-cut method compares reasonably well to the rigorous manual approach and the computer-generated model which also require only one performance test run.

►Both manual methods rely on a visual interpolation and extrapolation of the generic molecular sieve decline curves. The computer-generated approach provides much more consistent results.

►All the methods presented in these Tips of the Month rely on open-art technology. The molecular sieve vendors use proprietary methods specific to their manufacturing techniques. Consequently, the results of the approaches presented in these Tips of the Month should be used to generate trends as opposed to absolute values.

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►Site-specific factors will determine your unit’s decline curve. Conducting more than one performance test is highly recommended.

►Standby time offers a large degree of operating flexibility because the decline curves tend to level off; always try to build in standby time in any new molecular sieve design.

►Adsorption capacity is a function of the number of cycles, not calendar time.

►Install a good filter coalescer or filter separator upstream of your adsorption unit to keep the contaminants out of the system.To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing) and PF-4 (Oil Production and Processing Facilities) courses.

References1. Malino, H.M., http://www.jmcampbell.com/tip-of-the-month/2015/05/benefits-of-standby-time-in-adsorption-dehydration-process/, PetroSkills – John M. Campbell, 2015

2. Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9th Edition, 3nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2018.

3. Moshfeghian, M. http://www.jmcampbell.com/tip-of-the-month/2016/06/projecting-the-performance-of-adsorption-dehydration-process/, PetroSkills – John M. Campbell, 2016

4. GCAP 9.2.1 Software, PetroSkills – John M Campbell “Gas Conditioning and Processing Computer Program,” Editor Moshfeghian, M., PetroSkills, Katy, Texas, 2016.

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6060

IMPACT OF TEMPERATURE APPROACH OF THE HEAT EXCHANGERS

ON THE CAPEX AND OPEX OF A MECHANICAL REFRIGERATION PLANT

WITH MEG INJECTION

Published June 2019

Written by: Dr. Mahmood Moshfeghian

The design or specified minimum temperature approach of a heat exchanger has a significant effect on the total heat transfer area required. As a result, the specified temperature approach (TA) should be carefully considered in the heat exchanger specification process as this is one of the primary factors in heat exchanger capital cost. Depending upon heat exchanger service and type, there are typical economic minimum temperature approaches which have been determined from industry experience.Continuing the April 2019 [1] Tip of The Month (TOTM), this tip investigates the impact of the TA on the performance of a mechanical refrigeration plant with mono-ethylene glycol (EG or MEG) injection for hydrocarbon dew point (HCDP) control. Specifically, how the TA impacts the gas-gas heat exchanger, chiller duties

and the operation of the mechanical refrigeration system will be investigated and reported. In addition, the annualized CAPEX, Annual OPEX (Energy Cost) and annual total cost as a function of the gas-gas heat exchanger (HX) hot end TA will be reported. The details of a mechanical refrigeration plant with MEG injection and regeneration system are given in Chapters 6, 15 and 18 of the Gas Conditioning and Processing, Volumes 1 and 2 [2, 3], respectively.Figure 1 presents the process flow diagrams for a typical HCDP control plant using mechanical refrigeration with MEG injection system. This figure is similar to the Apr 2019 TOTM [1] which utilizes a flash tank economizer with two stages of compression. In this tip, all simulations were performed with UniSim Design R443 software [4] using the Peng-Robinson equation of state.

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Introduction (Extracted from Chapter 12 of Reference [3]):The surface area of a heat exchanger can be calculated by Equation 1.

(1)

Where:A = heat exchanger area in m2 [ft2]

Q = heat transfer rate in kW [Btu/hr]

Uo = overall heat transfer coefficient in W/ (m2 oC) [Btu/(hr-ft2-oF)]

ΔTeff = effective temperature difference oC [oF]Figure 2 presents a temperature profile as a function of heat transfer in a chiller and a gas-gas HX. In the chiller, the process fluid is partially condensing in the tube side, and pure propane is boiling in the shell side. In the gas-gas HX, the process fluid (stream 5) is partially condensing in the tube side and the gas stream 4 is warming (no phase change) in the shell side. The only temperatures that can be conveniently measured are at the inlet and outlet of the exchanger. The largest difference on one end of the heat exchanger will be referred to as ΔT1, the smaller ΔT2, which is also called the minimum approach.

Fig 1. Process flow diagrams for an HCDP plant using mechanical refrigeration with a flash tank economizer and MEG Injection system

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Fig 2. Effective Temperature Difference Schematics [2]

Equation 2 provides a simple method for estimating the effective temperature difference (ΔTeff) in the exchanger.

(2)

Where:LMTD = log mean temperature difference

ΔTeff = temperature difference corrected for heat exchanger configuration

F = TEMA MTD Correction Factor

ΔT1 = largest ΔT (at one end of the heat exchanger)

ΔT2 = smallest ΔT (at one end of the heat exchanger) Equation 2 is based on several assumptions. The primary one being that the heating and cooling curves (T vs Q) for the exchanger streams are linear. For multi-component fluids this is true when there is little or no phase change in the exchanger. If you apply this equation, always have a look at the cooling curves to verify this assumption is adequate the application.Notices in this equation that as ΔT2 (minimum approach) decreases, ΔTeff decreases and the required heat transfer area increases as shown in Equation 1. This increase can be significant as ΔTeff approaches zero. On the other hand, smaller values of ΔT2 in the gas-gas exchanger for a refrigeration plant decreases the utility costs (power and fuel) there is more “free cold energy” transferred in the heat exchanger. For this

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specific process application (gas-gas exchanger) both effects should be considered in heat exchanger specifications. Note that the minimum approach may occur at the “hot” end or “cold” end of the exchanger depending on the application. The minimum approach may also occur internally inside the exchanger. Typical economic minimum approaches for various heat exchangers and applications are presented in Table 1. For compact exchangers, smaller minimum approaches are usually economically justifiable. Compact heat exchangers have a greater surface area per volume than shell and tube heat exchangers. In addition, they have higher fluid velocities, which result in greater overall heat transfer coefficients.

Case Study:Let’s consider the same case presented in Apr 2019 TOTM [1] for a rich gas with the compositions and conditions presented in Table 2 [1]. Based on the reported molecular weight and relative density for the

C7+ fraction, Table 3 presents the estimated normal boiling point (NBP), critical properties and acentric factor which, are needed by the equation of state. The objective is to meet a hydrocarbon dew point specification of -20 °C (-4 °F) at about 4000 kPa (580 psia) for the sales gas by removing heat in the “Gas/Gas” heat exchanger (HX) with a hot end TA of 5°C (9°F) and in a propane chiller, 5 °C (-4 °F) TA, and rejecting it to the environment by a propane condenser (AC-100) at 37.8°C (100°F). Pure propane is used as the working fluid in the simulation. The pressure drops in the “Gas/Gas” HX and the propane chiller are assumed to be 34.5 kPa (5 psi).

Table 2. Rich feed gas compositions and conditions

Table 1. Typical Heat Exchanger Approach Ranges [3]

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The feed gas is flashed in the “Inlet Separator” at 30 °C (86 °F) and 4000 kPa (580 psia) to remove any condensate. The “Inlet Separator” vapor (stream 2) is saturated with water by the “Saturate -100” to form stream “2 Wet” upstream of mixing with MEG hydrate inhibitor, stream “EG1” and the recycle stream “18A” from the deethanizer overhead vapor (located at the right-hand side of Fig. 1).The estimated hydrate formation temperature (HFT) of streams 5 and 7 is 14.7 °C (58.4 °F). The hydrate inhibitor is injected at the inlet of “Gas/Gas” HX by stream “EG1” and at the inlet of the “Chiller” by stream “EG2”. Stream “5” cools to about -8 °C (17.6 °F) and stream “7” cools down to the specified temperature of -20 °C (-4 °F), which is below the HFT of 14.7 °C (58.4 °F). The injection rates of streams “EG1” and “EG2” for 80-weight % lean MEG and water solution are estimated manually or by the Adjust tool of UniSim. A design margin of 1 °C (1.8 °F) HFT below the cold temperature for streams “5” and “7” were assumed.Assuming a temperature approach of 5°C (9°F) and a 6.9 kPa (1 psi) pressure drop in the propane chiller “Chiller” shell side, the pressure of saturated propane vapor leaving the chiller is 203.3 kPa (29.5 psia), and at a temperature of -25°C (-13°F). Assuming no frictional losses in the suction line to the propane

compressors “K-101” and “K-102”, the resulting suction pressure is 203.3 kPa (29.5 psia). The condensing propane pressure at the specified condenser temperature of 37.8 °C (100 °F) is 1303 kPa (189 psi). The condenser “AC-100” frictional losses, plus the frictional losses in the piping from the compressor discharge to the condenser were assumed to be 34.5 kPa (5 psi); therefore, the discharge pressure of compressor “K-102” is 1338 kPa (194 psia). The compressors interstage pressure was determined by equalizing the power for “K-101” and “K-102”. The compressors adiabatic efficiency was assumed to be 75%.The cold Stream 7 is flashed in the 3-phase separator “V-102” at -20 °C (-°4F) and 3931 kPa (570 psia). The vapor stream “4” from this cold separator is used to cool down the incoming warm feed gas in the “Gas/Gas” HX. The heavy liquid stream “8B” (rich MEG solution) from the cold separator is regenerated in the regeneration unit (not shown in Fig. 1) and the lean 80 weight % MEG is recycled and used in streams “EG1” and “EG2”. The cold NGL stream “8” (light liquid phase) from the cold separator, “V-102”, is combined with the plant “Inlet Separator” condensate (stream “3”) in the mixer “Mix-101” to form stream “9” at about 5 °C (41 °F) and 3945 kPa (572.2 psia). To prepare the liquid to be fed to the deethanizer,

Table 3. Estimated C7+ properties [4]

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the process specification is to raise the temperature of the NGL product stream “9A” from about -4°C (25°F) and 1535 kPa (222.6 psia) to 20 °C (68 °F) and 1500 kPa (217.6 psia) in “E-102” HX. The process duty and the temperature of the NGL product stream is set by the deethanizer process requirements. The pressure drops in “E-102” HX is 35 kPa (5 psi).

Deethanizer Specifications and Performance:ike the Apr 2019 TOTM [1], the deethanizer column specifications are:

► To recover 90 mole percent of propane of the feed in the bottom product and

►. Ethane to propane mole ratio equal to 5 % in the bottom’s product

►. Top and bottom pressures are 1450 and 1500 kPa (210.3 and 217.6 psia); respectively

► Number of theoretical stages 12 plus the condenser and reboiler (determined by the material balance and column shortcut calculations)

The deethanizer simulation results are summarized in Table 4.

Table 4. Summary of deethanizer key design parameters

Impact of Gas/Gas HX Hot End Temperature Approach:The Gas/Gas HX utilizes the cold temperature of stream 4 at -20 °C (-°4F) to cool down stream 2A. The specified temperature approach (TA) sets the sales gas temperature, which is equal to stream 2A temperature minus the TA. Decreasing the hot end TA increases the heat duty of the Gas/Gas HX, which decreases the stream 5 temperature and reduces the required chiller duty. Typically, the Gas/Gas HX removes about 70% of the total required cooling duty to meet the specified sales gas dewpoint temperature. Therefore, the chiller duty decreases resulting in lower compressor power, decreased

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propane refrigerant circulation rate, and reduced propane refrigerant condenser duty requirements. This tip investigates the impact of the hot end TA for the Gas/Gas HX for a range of 1 to 11 °C (1.8 to 19.8 °F) while keeping the chiller TA constant at 5 °C (9 °F).Figure 3 presents the impact of TA on the MEG injection rate of streams EG1 and EG2 upstream of Gas/Gas HX and the chiller, respectively. As the hot end TA increases, stream 5 temperature increases but the stream

7 temperature remains constant at the set value; therefore, the HFT depression of stream 5 (d = HFT minus Stream 5 T) decreases and the required MEG injection rate of stream EG1 decreases. However, the MEG injection rate of stream EG2 increases because the total HFT depression temperature (d = HFT minus Stream 7 T) is constant; therefore, the total MEG injection (EG1 + EG2) rate stays the same. The calculated EG1 and EG2 were summed up and presented in Figure 3.

Fig 3. Impact of the temperature approach on the MEG injection rate upstream of Gas/Gas HX (EG1) and chiller (EG2)

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Increasing TA causes a decrease in the Gas/Gas-gas HX duty, which results in an increase in the chiller duty because the total cooling duty is constant. As the chiller duty decreases, the compressor power and condenser duty decreases, too. Figures 4 A and B illustrate the impact of TA on the compressor power, Gas/Gas HX, Chiller, and condenser duty in SI and FPS system of units, respectively. In order to calculate the required surface area of the heat exchangers a design factor of 1.25 and the following overall heat transfer coefficients in W/m2-°C (Btu/hr-ft2-°F) were assumed: Gas/Gas HX = 283 (50), Chiller = 444 (78), and Condenser = 460 (81).

Figure 5 presents the impact of TA on the required surface area of the Gas/Gas HX, chiller and the refrigerant condenser. This figure indicates that the Gas/Gas HX required area is much greater than the chiller and condenser area because it has greater heat duty. Note for TA cases of 1 and 2 °C (1.8 and 3.6 °F) the required surface areas of Gas/Gas HX are greater than 6000 m2 (64,590 ft2); therefore, two HXs of 50% of the required area should be used [5].

Fig 4A. Impact of the temperature approach on the Gas/Gas HX, chiller and condenser duty, and compressor power

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Fig 4B. Impact of the temperature approach on the Gas/Gas HX, chiller and condenser duty, and compressor power

Fig 5. Impact of the temperature approach on the Gas/Gas HX, chiller and condenser surface area 69

The procedure suggested by Hubbard [5] was used to perform the cost analysis. The CAPEX is the sum of the HXs and compressor costs. The HXs cost was estimated based on the size and surface area for the shell and tube types HXs. It is assumed that the maximum shell diameter is 1.97 m (6 ft) and the L/D was 10. For these dimensions, the maximum surface area per shell is about 6000 m2 (64,590 ft2). For TA cases of 1 and 2 oC (1.8 and 3.6 oF), the Gas/Gas HX with an area greater than 6000 m2 (64,590 ft2) two HXs with 50% required area were used. The compressor CAPEX was estimated based on the required power with a design factor of 1.25. Annualized

CAPEX cost was estimated based on 5 years recovery. The annual OPEX (energy cost) was estimated based on the compressor power requirement. The total cost is the sum of the annualized CAPEX and annual OPEX (energy cost). These costs are based on a design factor of 1.25. Because total MEG injection rate and deethanizer performance are independent of TA, they were excluded in the cost analysis.Figure 6 presents the impact of TA on the estimated annualized CAPEX, annual OPEX (energy cost), and the total annual cost. This figure indicates that the optimum TA for the Gas/Gas HX occurs at 3 °C (5.4 °F).

Fig 6. Impact of the temperature approach on the annual OPEX, annualized CAPEX and the total annual cost 70

Summary:This tip demonstrated the impact of temperature approach (TA) on the trade-off between CAPEX and OPEX for the mechanical refrigeration of a HCDP control plant. The heat exchanger costs are based on rough estimates; therefore, the economic evaluation presented does not represent the actual costs. The Gas/Gas HX areas are extremely large for TA cases of 1 and 2 oC (1.8 and 3.6 oF). It is unlikely that Shell and Tube exchangers of this size could even be manufactured. The Gas/Gas HX CAPEX for the TA cases of 1 and 2 oC (1.8 and 3.6 oF) assumed two exchangers of the 50 % required surface area. If a new mechanical dewpoint facility were built today, it is probably that the Gas/Gas HX and the chiller would almost certainly be a compact heat exchanger, such as a plate-fin or printed circuit. The compact heat exchanger designs have significantly greater surface area per volume than shell and tube heat exchangers which allows them to achieve smaller economic temperature approaches for a given capital cost.The minimum cost is at 3 °C (5.4 °F), which is where we must go to 2 parallel shells, which makes sense because here we have a step change in the configuration. When the temperature correction factor, F < 0.8, we usually add another shell in series, which may also give step changes. Typically these exchangers have 2 or 3 shells in series at the final configuration.The minimum cost is at 3 °C (5.4 °F), which is where we must

go to 2 parallel shells, which makes sense because here we have a step change in the configuration. When the temperature correction factor, F < 0.8, we usually add another shell in series, which may also give step changes. Typically these exchangers have 2 or 3 shells in series at the final configuration.To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing) and PF42 (Separation Equipment - Sizing and Selection) courses.

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References:1. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2019/04/Impact-of -liquid-carryover-on-performance-of-refrigeration-system-with-a-flash-tank-economizer/, PetroSkills -John M. Campbell Tip of the Month, April 2019.

2. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum

3. Campbell, J.M., “Gas Conditioning and Processing, Volume 2: The Equipment Modules,” 9th Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018.

4. UniSim Design R443, Build 19153, Honeywell International Inc., 2017.

5. Hubbard, R.A., Personal Communication, March 2019.

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