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.
Figure 1. Process flow diagrams for a HCDP plant using mechanical refrigeration with a flash tank economizer and MEG Injection system
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]. 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]
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 wasdetermined 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).
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
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 drybasis.
Figure 2. Impact liquid carryover on the sales gas hydrocarbon dewpoint temperature
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).
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 5 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.
Figure 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
Figure 4. Impact of liquid carryover on the MEG injection rate upstream of Gas/Gas HX (EG1) and chiller (EG2)
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.
Figure 5A. Impact of liquid carryover on the compressor power, Gas-Gas HX, chiller, and condenser duty
Figure 5B. Impact of liquid carryover on the compressor power, Gas/Gas HX, chiller, and condenser duty
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 alsoincreases. 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
SUMMARY:
The common practice to meet sales gas hydrocarbon dewpoint temperature under the condition of liquidcarry 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 ourG4 (Gas Conditioning and Processing), G5 (Practical Computer Simulation Applications in Gas Processing), and PF42 (Separation Equipment – Sizing and Selection) courses.
By: Dr. Mahmood Moshfeghian
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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.
