\u25baThe velocity profile through the mist extractor is poor, resulting in localized high velocities\/flooding.<\/p>\n
\u25baThe droplet sizes reaching the mist extractor are too small.<\/p>\n
\u25baThe entrained liquid load reaching the mist extractor is too high.<\/p>\n
\u25baThe mist extractor is damaged or plugged.<\/p>\n
\u25baLevel control and instrumentation malfunction or failure<\/p>\n
\u25baFoaming<\/p>\n
<\/p>\n
Continuing the\u00a0December<\/a>\u00a02005,\u00a0January<\/a>\u00a0and\u00a0February<\/a>\u00a02019 [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.<\/p>\n <\/p>\n 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\u00a0carry\u00a0in separation equipment are discussed in PetroSkills-John M. Campbell course titled, \u201cPF42 – Separation Equipment – Sizing and Selection.\u201d<\/p>\n <\/p>\n 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.<\/p>\n <\/p>\n <\/p>\n Figure 1.<\/strong>\u00a0Process flow diagrams for\u00a0<\/em>a HCDP<\/em>\u00a0plant using mechanical refrigeration with a flash tank economizer and MEG Injection system<\/em><\/p>\n<\/div>\n <\/p>\n <\/p>\n CASE STUDY:<\/strong><\/p>\n Let\u2019s 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 \u00a0-20 \u00b0C (-4 \u00b0F) at about 4000 kPa (580\u00a0psia) for the sales gas by removing heat in the \u201cGas\/Gas\u201d heat exchanger (HX) with a hot end approach temperature of 5\u00b0C (9\u00b0F) \u00a0and in a propane chiller, 5 \u00b0C (-4 \u00b0F) approach temperature, and rejecting it to the environment by a propane condenser (AC-100) at 37.8\u00b0C (100\u00b0F). Pure propane is used as the working fluid in the simulation. The pressure drops in the \u201cGas\/Gas\u201d HX and the propane chiller are assumed to be 34.5 kPa (5 psi).<\/p>\n <\/p>\n <\/p>\n Table 1.<\/strong>\u00a0Rich feed gas compositions and conditions<\/p>\n <\/p>\n <\/p>\n Table 2.<\/strong>\u00a0Estimated C7+ properties [4]<\/p>\n <\/p>\n <\/p>\n The feed gas is flashed in the \u201cInlet Separator\u201d at 30 \u00b0C (86 \u00b0F) and 4000 kPa (580\u00a0psia) to remove any condensate. The \u201cInlet Separator\u201d vapor (stream 2) is saturated with water by the \u201cSaturate -100\u201d to form stream \u201c2 Wet\u201d upstream of mixing with MEG hydrate inhibitor, stream \u201cEG1\u201d and the recycle stream \u201c18A\u201d from the\u00a0deethanizer\u00a0overhead vapor (located at the right-hand side of Fig. 1).<\/p>\n <\/p>\n The estimated hydrate formation temperature of streams \u201c2 Wet\u201d is 14.7 \u00b0C (58.4 \u00b0F). The hydrate inhibitor is injected at the inlet of \u201cGas\/Gas\u201d HX by stream \u201cEG1\u201d and at the inlet of the \u201cChiller\u201d by stream \u201cEG2\u201d. Stream \u201c5\u201d cools to about -8 \u00b0C (17.6 \u00b0F) and stream \u201c7\u201d cools down to the specified temperature of -20 \u00b0C (-4 \u00b0F) which are below the hydrate formation temperature (HFT) of 14.7 \u00b0C (58.4 \u00b0F). The injection rates of streams \u201cEG1\u201d and \u201cEG2\u201d for 80 weight % lean MEG and water solution are estimated by the Adjust tool of UniSim. A design margin of 1 \u00b0C (1.8 \u00b0F) HFT below the cold temperature for streams \u201c5\u201d and \u201c7\u201d were assumed.<\/p>\n <\/p>\n Assuming an approach temperature of 5\u00b0C (9\u00b0F) and a 6.9 kPa (1 psi) pressure drop in the propane chiller (\u201cChiller\u201d) shell side, the pressure of saturated propane vapor leaving the chiller is 203.3 kPa (29.5\u00a0psia), and at a temperature of -25\u00b0C (-13\u00b0F).\u00a0 Assuming no frictional losses in the suction line to the propane compressors \u201cK-101\u201d and \u201cK-102\u201d, the resulting suction pressure is 203.3 kPa (29.5\u00a0psia).<\/p>\n <\/p>\n The condensing propane pressure at the specified condenser temperature of 37.8 \u00b0C (100 \u00b0F) is 1303 kPa (189 psi). The condenser \u201cAC-100\u201d 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 \u201cK-102\u201d is 1338 kPa (194\u00a0psia). The compressors\u00a0inter stage\u00a0pressure\u00a0wasdetermined by equalizing the power for \u201cK-101\u201d and \u201cK-102.\u201d The compressors adiabatic efficiency was assumed to be 75%.<\/p>\n <\/p>\n The cold Stream 7 is flashed in the 3-phase separator \u201cV-102\u201d at -20 \u00b0C (-\u00b04F) and 3931 kPa (570\u00a0psia). The vapor stream \u201c4\u201d from this cold separator is used to cool down the incoming warm feed gas in the \u201cGas\/Gas\u201d HX. The heavy liquid stream \u201c8B\u201d (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 \u201cEG1\u201d and \u201cEG2\u201d. The cold NGL stream \u201c8\u201d (light liquid phase) from the cold separator, \u201cV-102\u201d, is combined with the plant \u201cInlet Separator\u201d condensate (stream \u201c3\u201d) in the mixer \u201cMix-101\u201d to form stream \u201c9\u201d at about 5 \u00b0C (41 \u00b0F) and 3945 kPa (572.2\u00a0psia). To prepare the liquid to be fed to the\u00a0deethanizer, the process specification is to raise the temperature of the NGL product stream \u201c9A\u201d from about -4\u00b0C (25\u00b0F) and 1535 kPa (222.6\u00a0psia) to 20 \u00b0C (68 \u00b0F) and 1500 kPa (217.6\u00a0psia) in \u201cE-102\u201d HX. The process duty and the temperature of the NGL product stream\u00a0is\u00a0set by the\u00a0deethanizer\u00a0process requirements. The pressure drops in \u201cE-102\u201d HX is 35 kPa (5 psi).<\/p>\n <\/p>\n <\/p>\n DEETHANIZER\u00a0SPECIFICATIONS and PERFORMANCE:<\/strong><\/p>\n Like the February 2019 TOTM [3], the\u00a0deethanizer\u00a0column specifications are:<\/p>\n A. To recover\u00a090 mole\u00a0percent of propane of the feed in the bottom product and<\/p>\n B. Ethane to propane mole ratio equal to 5 % in the bottoms product<\/p>\n C. Top and bottom pressures are 1450 and 1500 kPa (210.3 and 217.6\u00a0psia); respectively<\/p>\n D. Number of theoretical stages 12 plus the condenser and reboiler (determined by the material balance and column shortcut calculations)<\/p>\n The\u00a0deethanizer\u00a0simulation results are summarized in Table 3.<\/p>\n <\/p>\n <\/p>\n Table 3.<\/strong>\u00a0Summary of\u00a0deethanizer\u00a0key design parameters<\/p>\n <\/p>\n <\/p>\n IMPACT\u00a0OF LIQUID HYDROCARBON CARRYOVER:<\/strong><\/p>\n Separator \u201cV-102\u201d is a three-phase separator. Under ideal\u00a0condition\u00a0the vapor (stream 4) leaving the separator has no LCO and its dewpoint temperature is the same as the feed (stream 7) temperature.Typical\u00a0range of liquid\u00a0carry over\u00a0is 0.013\u20130.27 m3<\/sup>\u00a0liquid\/106<\/sup>\u00a0std m3<\/sup>\u00a0of gas (0.1\u20132 gallon of liquid\/MMscf) [5]. In\u00a0practice\u00a0due to the reasons listed in the preceding\u00a0section\u00a0the LCO can be even higher. In this tip, the impact of LCO was investigated for a range of 0 to 3 mole % of\u00a0liquid\u00a0in\u00a0light\u00a0liquid 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\u00a0separator\u00a0(\u201cV-102\u201d). This is possible if the mechanical refrigeration system is capable of handling a higher chilling load.<\/p>\n <\/p>\n 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\u00a0negligible\u00a0effect on the bubble curves. All phase envelopes are generated on\u00a0the drybasis.<\/p>\n <\/p>\n <\/p>\n Figure 2.<\/strong>\u00a0Impact liquid carryover on the sales gas hydrocarbon dewpoint temperature<\/em><\/p>\n <\/p>\n <\/p>\n 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 \u00b0C (-4 \u00b0F). For 3 mole % LCO, the sales gas dewpoint temperature is -14.4 \u00b0C (6.1 \u00b0F). To bring back the sales gas dewpoint temperature, the process gas (stream 7) should be cooled to -28.6 \u00b0C (-19.5 \u00b0F).<\/p>\n <\/p>\n 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.<\/p>\n <\/p>\n If there is\u00a0carryover\u00a0of the hydrocarbon phase, there is also a likelihood of carryover of the glycol phase.\u00a0 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.<\/p>\n <\/p>\n <\/p>\n Figure 3.<\/strong>\u00a0Impact of liquid carryover on the sales gas dewpoint temperature (solid line) and the required cold separator feed temperature (dashed line) to offset liquid carryover<\/em><\/p>\n <\/p>\n <\/p>\n Figure 4.<\/strong>\u00a0Impact of liquid carryover on the MEG injection rate upstream of Gas\/Gas HX (EG1) and chiller (EG2)<\/em><\/p>\n <\/p>\n 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.<\/p>\n <\/p>\n <\/p>\n Figure 5A.<\/strong>\u00a0Impact of liquid carryover on the compressor power, Gas-Gas HX, chiller, and condenser duty<\/em><\/p>\n <\/p>\n <\/p>\n Figure 5B.\u00a0<\/strong>Impact of liquid carryover on the compressor power, Gas\/Gas HX, chiller, and condenser duty<\/em><\/p>\n <\/p>\n Figure 6 presents the impact of LCO on\u00a0the liquid\u00a0propane 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\u00a0mole\u00a0basis. The extra propane liquid recovery is achieved by operating the chiller at a lower temperature which requires higher OPEX and CAPEX.<\/p>\n <\/p>\n As the chiller temperature is reduced, more ethane and methane end up in the\u00a0low temperature\u00a0separator (LTS), V-102, liquids.\u00a0 These need to get boiled out in the\u00a0deethanizer, so the duty of E-102 increases, reboiler and condenser duty of the\u00a0deethanizer\u00a0increase, and the recompression power in K-100 alsoincreases.\u00a0It may also be possible to flood the\u00a0deethanizer.<\/p>\n <\/p>\n The cold condenser on the\u00a0deethanizer\u00a0requires propane for cooling. These units are also designed with no condenser.\u00a0The cold LTS (V-102) liquid is used as reflux, and the liquids from the inlet separator are introduced lower in the column.\u00a0V-100 is not required then. E-102 is usually a feed\/bottoms heat exchanger. Overhead of the\u00a0deethanizer\u00a0can likely go to the sales gas and may not have to be recycled.\u00a0 It really should not contain anything heavier than propane, heavy key (HK).<\/p>\n <\/p>\n Table 4 presents the impact of LCO on the key equipment incremental capacity\u00a0requirement\u00a0to 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\u00a0condition\u00a0it may require plant shutdown for\u00a0trouble shooting\u00a0to reduce the LCO.<\/p>\n <\/p>\n Table 4.<\/strong>\u00a0Estimate of equipment incremental capacity\u00a0requirement\u00a0to handle liquid carryover<\/p>\n <\/p>\n <\/p>\n SUMMARY:<\/strong><\/p>\n The common practice to meet sales gas hydrocarbon dewpoint temperature under the condition of liquidcarry\u00a0is 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\u00a0and\u00a0the refrigeration requirement.<\/p>\n <\/p>\n 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\u00a0operation\u00a0conditions 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\u00a0and\u00a0liquid gravity separation sections.<\/p>\n <\/p>\n To learn more about similar cases and how to minimize operational problems, we suggest attending ourG4 (Gas Conditioning and Processing)<\/a>,<\/strong>\u00a0G5 (Practical Computer Simulation Applications in Gas Processing)<\/a>,\u00a0<\/strong>and\u00a0PF42 (Separation Equipment – Sizing and Selection)<\/a><\/strong>\u00a0courses.<\/p>\n By: Dr. Mahmood Moshfeghian<\/em><\/p>\n <\/p>\n Sign up to receive monthly Tip of the Month emails<\/em>!<\/em><\/p>\n![\"\"](\"https:\/\/i0.wp.com\/www.jmcampbell.com\/tip-of-the-month\/wp-content\/uploads\/2019\/04\/fig-1-e1556807148515.png?resize=600%2C248\")
<\/a><\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/t-1.png?ssl=1\")
<\/p>\n![\"\"](\"https:\/\/i0.wp.com\/www.jmcampbell.com\/tip-of-the-month\/wp-content\/uploads\/2019\/04\/t-2-e1556807220857.png?resize=600%2C63\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/t-3.png?ssl=1\")
<\/p>\n![\"\"](\"https:\/\/i0.wp.com\/www.jmcampbell.com\/tip-of-the-month\/wp-content\/uploads\/2019\/04\/fig-2-e1556807285267.png?resize=600%2C374\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/fig-3.png?ssl=1\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/fig-4.png?ssl=1\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/fig-5a.png?ssl=1\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/fig-5b.png?ssl=1\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/images\/apr19-fac\/t-4.png?ssl=1\")
<\/p>\n![](\"https:\/\/i0.wp.com\/www.petroskills.com\/logos\/ps-jmc_rgb-150.png?ssl=1\")
<\/p>\n
\n