Benzene, toluene, ethylbenzene, and xylene are members of the aromatics hydrocarbon family group, often referred to as BTEX. These aromatic compounds are also belonged to the broader category of Hazardous Air Pollutants (HAPs). Benzene is a known carcinogen and has also been shown to cause blood disorders and to impact the central nervous system and the reproductive system.  Toluene may affect the reproductive and central nervous systems.  Ethylbenzene may have respiratory and neurological effects [1]. BTEX can be present in many natural gas streams and are partially absorbed by the solvent in glycol dehydration and amine sweetening units.

In gas treating service, methyl diethanolamine (MDEA) will absorb limited quantities of BTEX from the gas. Based on literature data, predicted absorption levels for BTEX components vary from 5 to 30% [2]. Absorption is favored at lower temperatures, higher MDEA concentrations and circulation rates. The bulk of absorbed BTEX is separated from the MDEA in the regeneration unit and leaves the system in the regenerator overhead stream which requires further treatment before being vented to atmosphere.  The amount of adsorption of the BTEX is required to be known to determine the proper treating method for the overhead acid gases leaving the regenerator.

The emission of BTEX components from glycol dehydration is also regulated in most countries.  In the U.S., benzene emissions are limited to 1 ton/year (900 kg/year).  Mitigation of BTEX emissions is an important component in the design of a dehydration systems.  Correctly estimating the quantity of absorbed BTEX and understanding the factors that affect absorption levels is critical to ensure the proper mitigation methods are provided to meet the required emission limits.

The GPA Midstream research report RR-242 [3], is an extension of several previous GPA Midstream research projects looking at the solubility of hydrocarbons in loaded and unloaded amine solutions.  The previous research included research reports RR-180, 185, 195, and RR-220.

The previous projects have concentrated only on two amine systems, MDEA and DGA, both loaded and unloaded with several model hydrocarbons. RR-242 expands the current base of research data to other amines (including DEA, MEA, and a MDEA/piperazine blend), as well as measures the influence of CO₂ and H₂S (so called loaded amines).

Accurate hydrocarbon solubility data of RR-242 enables the development of new equation of state correlations that can be applied to the simulation of amine units (Bullin and Brown [4]). The data can be used to optimize the design and operation of amine units in which these hydrocarbons are present in the feed gas. The data will provide a basis for accurately predicting the distribution of the heavier hydrocarbons between the treated gas, the amine flash gas, and the acid gas streams. The data will provide a basis for accurately predicting emissions of these hydrocarbons from the amine unit to aid in the design and operation of these units (Moshfeghian and Hubbard [5]).

This tip demonstrates applications of the models developed based on the experimental VLE and VLLE data measurements presented in RR-242 [3]. The models can be used for estimation of solubilities of selected hydrocarbons in pure water, unloaded amine solution, loaded amine solution with CO₂ and/or H₂S. Because all selected hydrocarbons, both soluble and absorbed, in the contactor will be removed from the solution in the flash drum and regenerator (see Figure 1), the emissions will be essentially equal to the amount absorbed.

Figure 1- Process flow diagram of MDEA acid gas recovery unit [2]

This tip will focus on the following design variables:

– Solubility of selected hydrocarbons in pure water, Equation 1, and Figure 2.

– GPA Midstream RR-242 Proposed Model

– Relative solubility of benzene, toluene, and ethyl benzene (EBenzene), Figures 3, 4, and 5.

–  Relative solubility of benzene, toluene, and ethyl benzene (EBenzene), Figure 6.

– Solubility of benzene, toluene, and Ebenzene in 50 wt % MDEA solution, Figure 7, and Table 1.

Solubility of selected hydrocarbons in pure water:

Solubility of selected hydrocarbons in pure water can be estimated using Equation 1 [3] or Figure 2.

           (1)

Where T is the temperature in K and the correlation parameters, E, F and T_min, are shown in Table 1.

Table 1- Equation 1 parameters, E, F and T_min [3].

Figure 2 was generated using Equation 1 and its parameters in Table 1. The solubility of these selected hydrocarbons in pure water will be used to estimate their solubilities in amine solution in the proceeding sections.

Figure 2. Solubility of selected hydrocarbons in pure water [3].

Estimation of solubility of hydrocarbons in amine solution by GPA Midstream RR-242 Model [3]

Equation 2 with its corresponding parameters will be used to estimate the solubility of selected hydrocarbons in unloaded and loaded amine solutions.

     (2)

Where:

The experimental measurement data of relative solubility of benzene, toluene, and ethyl benzene (EBenzene) in 50 wt % MDEA solutions are shown in Figures 3, 4, and 5. In addition the estimated solubility data by Equation 2 is superimposed on these three figures.

Figure 3 – Relative solubility of Benzene in 50 wt % (13.1 mol-%) MDEA-Water [Fig 16 of GPA RR 242].

Figure 4 – Relative solubility of Toluene in H₂S or CO₂ loaded 50 wt % (13.1 mol-%) MDEA-Water [Fig 18 -GPA RR 242].

Figure 5 – Relative solubility of Ethylbenzene in CO₂ loaded 50 wt % (13.1 mol-%) MDEA-Water [Fig 20-GPA RR 242].

Tables 2 and 3 indicate that estimated solubility by the model agree well with the experimental measurements.

Table 2 – Solubility of Toluene in CO₂-loaded aqueous MDEA (13.13 mol-%)/water [3].

n – number of analyzed samples, x*–average, σ – standard deviation, x/x₀ – soluble loaded/soluble. non-loaded

Table 3 – Solubility of Toluene in H₂S-loaded aqueous MDEA (13.13 mol-%)/water [3].

APD = Absolute percent deviation, AAPD = Average Absolute percent deviation,

Example 1

Determine the solubility of benzene in loaded 13.1 mol% (50 wt%) MDEA in water solution in terms of scf of gas/gal of solvent (std m³ of gas/m³ of solvent) at 60 °C (140 °F), 333 K (600 °R).

Rich amine solution acid gas loading,  = 0.4 mol acid gases/ mole MDEA, MDEA MW =119.17

50 wt % MDEA solution density =1017.3 kg/m³ (8.49 lbm/gal)

Solution

From Fig3, for  = 0.4 mol acid gases/ mole MDEA, relative solubility, x/x₀= 0.53

Benzene solubility in un-loaded MDEA, x₀ = 0.00464

Figure 3 – Relative solubility of Benzene in 50 wt % (13.1 mol-%) MDEA-Water [Fig 16 of GPA RR 242].

Solution – SI

X = x₀(x/x₀) = (0.00464) (0.53kmol Ben/kmol MDEA Sol) (kmol MDEA Sol/0.131kmol MDEA)

   = (0.0187 kmol Ben/kmol MDEA) (23.64 std m³/kmol Ben) (kmol MDEA/119.21 kg)

    = (0.0037 std m³ Ben/kg of MDEA) (kg MDEA/2 kg MDEA + Water Sol) (1017.3 kg/m³ MDEA Sol)

    = 1.88 std m³ of Ben/m³ of MDEA + water solution

Or X = (770) (x₀) (x/x₀) std m³ of Ben/m³ of MDEA + water solution

Solution – FPS

X = x₀(x/x₀) = (0.00464) (0.53 lbmol Ben/lbmol MDEA Sol) (lbmol MDEA Sol/0.131lbmol MDEA)

   = (0.0187 lbmol Ben/lbmol MDEA) (379.5 scf/lbmol Ben) (lbmol MDEA/119.21 lbm)

    = (0.0598 scf Ben/lbm of MDEA) (lbm MDEA/2 lbm MDEA + Water Sol) (8.49 lbm/gallon MDEA Sol)

    = 0.25 scf of Ben/gallon of MDEA + water solution

Or X = (103) (x₀) (x/x₀) scf of Ben/gallon of MDEA + water solution

For simplicity and easiness of reading the figures, using the model (Equation 2) similar charts like Figures 6 and 7 or Table 4 can be generated.

Figure 6 – Relative solubility of benzene (Fig 3), toluene (Fig 4), and Ethylbenzene (Fig 5) in loaded 50 wt % (13.1 mol-%) MDEA-Water

Figure 7- Solubility of benzene, toluene, and ethyl benzene in unloaded 50 wt % MDEA solution

Table 4- Solubility of benzene, toluene, and ethyl benzene in unloaded 50 wt % MDEA solution

Figure 8 shows the influence of amine solution concentration and acid gas loading on solubility of toluene in 50 wt% and 25 wt% MDEA solution at 100 °F (38 °C)

Figure 8- solubility of toluene in 50 wt% and 25 wt% MDEA solution at 100 °F (38 °C)

From the RR-242 developed model based on the experimental results; several conclusions can be drawn.

1. The solubility of the hydrocarbons depends strongly on the amine concentration in the solvent, with the solubility increasing exponentially with increasing molar concentration of the amine in the aqueous solution (Figure 8).

2. The solubility of the hydrocarbons increases with temperature (Figure 7 and Table 4).

3. The solubility of the hydrocarbons depends strongly on the acid gas loading (decreased solubility at higher loading, Figure 8.) 

4. The solubility for CO₂ loaded amines are found to be similar to the solubility for H₂S loaded amines.

5. Finally, the ratio of hydrocarbon solubility of the loaded solvent to the unloaded solvent is found to be similar at a given acid gas loading.

Summary

Figures 6 and 7 present a simple tool for quick estimation of solubility of BETX compounds in MDEA gas treating process. For the example considered in this tip, the estimated solubility for each of BETX compounds matched well with the experimental results.

To learn more about similar cases and how to minimize operational problems, we suggest attending ourG4 (Gas Conditioning and Processing),G5 (Advanced Applications inGas Processing), http://www.jmcampbell.com/co2-surface-facilities-pf81.php andPF49 (Troubleshooting Oil & Gas Processing Facilities),courses.

By: Mahmood Moshfeghian, Ph.D.

and Kindra Snow-McGregor, P.E.

Reference:

1. http://www.earthworksaction.org/BTEX.cfm, 2011.

2. Campbell, J. M. “Gas conditioning and processing, Volume 2: The Equipment Modules,” 9^th Edition, 2^nd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.

3. Uusi-Kyyny, P., Pakkanen, M., Richon, D., Ionita, S., Ogunrobo, E., Alopaeus, V., RR-242, “Solubility of Hydrocarbons in Amine Treating Solutions”, GPA Midstream Association, Tulsa, OK, 2019.

4. Bullin, J. A., Brown, W.G., Hydrocarbons and BTEX Pickup and Control from Amine Systems”, 83^rd  Gas Processors Association Annual Convention, Mar. 2004.

5. Moshfeghian, M. and R.A. Hubbard, “Quick Estimation of Absorption of Aromatics Compounds (BTEX) in TEG Dehydration Process”, 3^rd International Gas Processing Symposium, March 5-7, Doha, Qatar, 2012.

Introduction

Following on the previous two TOTMs [1, 2] on Nord Stream long distance pipeline for natural gas transmission from Russia to Europe, this TOTM discusses the application of various long distance gas transmission correlations/equations that are available to determine the maximum gas capacity of a long-distance pipeline. In addition, calculations can be done to estimate the line packed gas volume; demonstrating that a long-distance pipeline can be used as a gas storage facility as well.

Case Study Data

In the previous TOTMs, we discussed various parameters that are available in the public domain [3] and presented in Table 1. Nord Stream 1 has 2 parallel gas pipelines, each pipeline capable of transporting gas from Russia to Germany at a rate of 75.34 million Std m³/day (2660.6 MMSCFD)

Table 1. Pipeline Specifications in SI and FPS Units

As the composition of the gas transported in the NS1 pipelines has not published, it was assumed for this study that the Norwegian gas, shown in Table 2, that is transported to the European Continental Shelf would be a good representation of the gas in the NS1 pipeline.

Table 2. Composition of the feed gas

Based on the above data, the physical and flow properties values were calculated and presented in Tables 3 and 4 in SI and FPS units, respectively.

Table 3. Calculated results for gas at a rate of 75.34 million Std m³/day

Table 4. Calculated results for gas at a rate of 2660.6 MMSCFD

In this study, two standard equations derived from the basic gas flow equation reflecting the average (mean) conditions have been used to estimate the NS1 pipeline capacity. These relationships are the BASIC and AGA (American Gas Association) equations containing a friction factor “f”. These equations are [4]:

The AGA equation estimates the friction factor under two conditions, “partially turbulent” and “fully turbulent”. The “partially turbulent” correlation is the Colebrook-White equation for smooth pipe. The smooth pipe assumption is corrected using a drag factor, F_f, which accounts for several pipeline characteristics: pipe wall condition, welds, changes in the direction of gas flow, isolation valves, etc. In the “fully turbulent” region the friction factor is independent of Reynolds number and depends only on the dimensionless roughness of the pipe.

These equations are available in GCAP software [5] and the screenshot of GCAP results are presented in Figures A1 – A5 in Appendix A.

Table 5. Calculated results for gas at a rate of 75.34 million Std m³/day

Table 6. Calculated results for gas at a rate of 2660.6 MMSCFD

The BASIC equation estimated gas flow of 84.07 million Std m³ (2969 MMscfd) of gas assuming 100 % efficiency. The BASIC equation estimated gas flow of 75.63 million Std m³ (2671 MMscfd) assuming 90 % efficiency.

The AGA equation estimated a gas flow of 84.59 million Std m³ of gas (2987 MMscfd) assuming 100% efficiency.

The published capacity of NS1 pipeline is 75.34 million Std m³ of gas [3].

A design flow rate check was also initiated by using PROMAX [6] that uses the single-phase regimes and single-phase basic equation applying Colebrook friction factor. This methodology gave a flow rate of 83.1 million Std m³ of gas.

The differences between the estimated and actual are 0.4% for 90% Basic, 11.6% for 100% BASIC, 12.30% for AGA and 10.3% with PROMAX. Considering that the gas composition is not known for NS1, these differences are reasonable. 

Pipe Wall Thickness

NS1 pipeline is varying in operating pressure, thus varying pipeline thickness. Using Barlow’s formula, a check on the wall thickness was made. Barlow’s formula relates the internal pressure that a given pipe can withstand as a function of its dimensions and strength of its material.

Where: 

 t = pipe wall thickness, mm (in)

P = 1.1 x design pressure, MPag (psig)

S = allowable stress, 483 MPag (70,000 psi) for grade x70 steel

D = outside diameter, mm (in)

In the Barlow’s formula, there is no provision of corrosion allowance. If we include this in the Barlow’s formula and take 3 mm as corrosion allowance (ca), then

Using Barlow’s equation, the pipeline wall thickness for different segments were calculated and presented in Table 7.

Table 7. Estimated pipeline wall thickness for different segments

Table 7 shows calculated pipeline wall thickness (t) with a corrosion allowance, ca. This information reveals a good agreement with the published data is reached using Barlow’s formula.

Line Packing

The NS1 gas pipeline can also be used to store gas when not used for gas transport. The methodology to determine how many standard cubic meters of gas can be held is given below.

Inventory (V) of gas pipeline can be calculated with:

Taking the internal diameter of the pipeline as 1.153 m (3.78 ft) and pipeline length of 1 224 000 m (4 015 748 ft), the pipeline volume is 1 277 997 m³ (12 578 711 ft³).

Then using the real gas law:

PV = nZRT

or molar volume

Where P is the average pressure between inlet (22000 kPa or 3190 psia) and outlet (10600 kPa or 1537 psia) pressures amounting to 16300 kPa (23663 psia), T is the average temperature in the pipeline (5 ⁰C=41°F or 278 K=500.4 °R), Z is the average compressibility of 0.75 and R is the universal gas constant (8.314 kPa.m³/kmol. K or 10.732 psia.ft³/lbmol-°R).

              n = 12 017 147 kmol (26 497 809)

Now 1 kmol = 23.64 Std m³ and 1 lbmol = 379.5 scf

Then the amount of standard cubic meters of gas that can be line-packed is 284.1 million Std m³ (10.6×10⁹ scf).

This is the amount of gas that can be stored in the pipeline when it is not operating in the transportation phase.

Concluding Remarks

The methodology used above demonstrates how standard long distance pipeline flow correlations/equations (BASIC and AGA equations) can be used in evaluating the design flow rates of an installed or of an operating pipeline. It was noted that the Basic equation with 90%-line efficiency matched the published design capacity very closely.

Pipe wall thickness calculation methodology using Barlow’s formula also shows good agreement with that of published data.

In addition, using the real gas law, it can be demonstrated that considerable amount of gas can be line-packed when the pipeline is in non-transportation mode.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G4 (Gas Conditioning and Processing), P81 (CO₂ Surface Facilities), and PF4 (Oil Production and Processing Facilities) courses.

References:

1. Moshfeghian, M., Rajani, J., and Snow-McGregor, K., “Transportation of Natural Gas in Dense Phase – Nord Stream”, PetroSkills TIP OF THE MONTH, April 2022.

2. Langer, J.F, Snow-McGregor, K, and Rajani, J., “Part 2: Nord Stream Pipelines – Multiple Parallel Paths to Succes or Failure?”, PetroSkills TIP OF THE MONTH, April 2022.

3. Beaubouef, B., “Nord stream completes the world’s longest subsea pipeline,” Offshore, P30, December 2011.

4. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^th Edition, 3^rd Printing, Editors Hubbard, R., and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

5. GCAP 10.2.1, Gas Conditioning and Processing, PetroSkills/Campbell, Tulsa, Oklahoma, 2022.

Appendix A: GCAP Results

Figure A1. Calorimetric Values

Figure A2. Average gas density by GCAP-Option 3C, SRK EOS

Figure A3. Estimated design pipeline capacity by basic Equation with 90 % efficiency

Figure A4. Estimated design pipeline capacity by basic Equation with 100 % efficiency

Figure A5. Estimated design pipeline capacity by AGA Equation

I. Introduction: Gas Lift Operational Fundamentals
Part 1 of this Series on Gas Lift History and Basic Well Parameters focuses on the primary “state of affairs” of Gas Lift operations in the USA. A discussion was presented related to a candidate Gas Lift well’s completion design that included a typical Casing/Tubing sizing sequence. The function of the production tubing gas lift Mandrels in starting a “kick – off” procedure in a candidate well were discussed. Types of Mandrel Gas Lift Valves were discussed, along with a discussion of the Single Gas Lift Valve (with its appropriate orifice size) employed as the final receptor of the injected casing gas.

II. Operational Fundamentals for the Performance of a Gas Lift Well, Related to Choke Flow, Single Phase Gas, and Multiphase Flowing Gradients
Part 2 of this Series reviewed energy and mass balances as related to a candidate Gas Lift Well’s flowing gradient. Energy balance equations provided the proper data to simulate both the annular flow in a casing/tubing configuration, as well as for choke (orifice) performance. Data from Industry standards provided the injected Casing/Tubing/Liner dimensions to calculate Effective Areas, as well as Effective Diameters for flow of the injected casing gas to the Production Tubing Gas Lift Valve at a given depth. The Thornhill–Craver equation provided the choke (orifice) performance data for the Gradient curves presented in Appendix A and B. The Pressure versus Gas Rate curves apply to orifice sizing but are only an estimate for gas lift valves since the valve stem in the seat reduces flow area.

Part 3 will review procedures for identifying, selecting, and optimizing technical as well as field operations for a gas lift well. Section IIIA reviews the gas lift well candidate related to gas content in the reservoir fluid and a choice of gas lift or pumping; Section IIIB discusses the well completion related to dimensional and clearance considerations and gas lift facility requirements; Section IIIC has guides for kicking off a well and avoiding erosion cutting of the unloading valves; Section IIID provides the procedure to optimize the well once it has kicked off and is operating in the production system.

Section IIIA Gas Lift Well Candidates
Reservoir conditions are primary drivers in choosing artificial lift. Gas content is key since gas lift supplements existing gas in the reservoir fluid and high content reduces the gas lift contribution. However, pumps are adversely affected by gas content in the reservoir fluid, leading to lower effectiveness and frequent failures. Production rate is a consideration that affects tubular size in gas lift, but changes pump choice with rod/beam pump for lower rates and electric submersible pump for higher rates. Finally, sand production from a sandstone reservoir or frac sand from a horizontal shale well have a detrimental effect on pump operation but a lesser effect on gas lift.

The fundamental relationship of gas lift and pumping with the reservoir is explained. Figure 1 shows a gas lift schematic on the left and the pressure‐rate behavior of flow from the reservoir (inflow) and up the tubing string (outflow) on the right. The gas lift well has gas entry through the tubing‐casing annulus to the operating valve or orifice, mixing with reservoir gas, oil, and water to reduce the composite density of the fluid, flowing to the wellhead and on to the separator. The chart at right shows the pressure versus rate for the inflow from the reservoir based on a PI = 1 bbl/d per psi and PR = 3000 psig. Each well is tested to obtain its productivity index (PI, related to IPR, Inflow Performance Relationship) and reservoir pressure (PR). This data is coupled with a nodal analysis program (PetroSkills uses SNAP from Tom Nations) to evaluate different tubing sizes and select the most appropriate. The other curves on the chart are multiphase flow results for 2 7/8” tubing outflow (called Vertical Lift Performance VLP, J‐Curve, Tubing Intake, Outflow) for natural flow and for gas lift. The intersection of the inflow curve and the outflow curve indicates a point of stable operation. The chart shows the natural flow curve (of higher density and pressure compared to gas lift) intersecting the inflow line at 400 stb/d whereas the gas lift curve (of lower density and lower bottomhole flowing pressure) intersecting at 1000 stb/d. The gas lift outflow curve is for a specific injection gas to liquid ratio (IGLR, scf/bbl) and although more injection gas would continue to decrease density, friction increase more than offsets the density reduction which is shown by the third curve in red, representing excessive IGLR that reduces production rate. IGLR plus formation gas to liquid ratio (FGLR) gives total gas to liquid ratio (TGLR). FGLR is a function of Solution Gas to Oil Ratio, Rs (scf/stb) and Oil Formation Volume Factor, Bo (bbl/stb). These two oil properties affect the quantity of “free flowing gas” which increases from the bottom of the bore to the wellhead as pressure and temperature reduce. Notice that all “outflow” curves decrease in bottomhole flowing pressure (negative slope) where density is the governing factor, and then increase (positive slope) as fluid velocity and friction become dominant.

Figure 1 – Gas Lift schematic and chart of pressure‐rate behavior (1)

Pumping is illustrated similarly to gas lift. Figure 2 shows a pump schematic on the left and on the right the pressure‐rate behavior of flow from the reservoir to the pump (inflow) and from the pump up the tubing string (outflow). A pumped well typically has no packer. Gas breaks out as reservoir fluids enter the casing and flows up the annular space to the wellhead where it is again mixed with produced fluids that are pumped up the tubing. The chart at right shows the pressure versus rate for the inflow from the reservoir based on a PI = 1 bbl/d per psi and PR = 3000 psig and represents flow into the pump suction. The other curve on the chart is multiphase outflow inside 2 7/8” tubing for natural flow and represents discharge from the pump. The chart shows the natural flow curve intersecting the inflow line at 400 stb/d, but to achieve 1000 stb/d with artificial lift, the fluid must be pumped from the intake pressure to the outflow pressure. The work energy put into the pump can be estimated from the pressure difference (Outflow pressure – Inflow Pressure), rate, and fluid density. Given more input power, a larger pump could continue to increase rate compared to gas lift, which is limited by friction.

Figure 2 – Pump schematic and chart of pressure‐rate behavior (1)

Nodal analysis simulation of gas lift and pumping is based on well testing to obtain reservoir inflow and to sample the reservoir fluids. This simulation permits evaluation of tubing sizes (and corresponding casing/liner sizes), changes in operation as reservoir conditions change, and the “best” amount of gas lift gas.

Section IIIB Gas Lift Well Dimensional and Gas Facility Considerations
Figure 3 is the gas lift well completion schematic used in Part 1 and 2. The arrangement of tubing, casing, and gas lift valves are shown, and distinction is made between unloading (kick off) valves and a bottom orifice for continuous injection.

Figure 3 – Gas lift well schematic and surface facility (2)

Prior to drilling, a tubing/casing size evaluation is conducted using exploration or prior well data as estimates for PI and reservoir pressure. If the reservoir has the rate capacity and engineering concurs that higher rates will not damage the reservoir rock nor preclude reserves recovery, then a larger size tubing and corresponding casing can be recommended. Once tubing/casing sizes are selected, a detailed dimensional analysis is required. The gas lift mandrels and associated valves are eccentric (both conventional tubing retrievable and side pocket with wireline retrievable valves) and clearance between the casing and tubing must be assured. The inside drift diameter of casing must be compared to tubing coupling outside diameter and to gas lift mandrel major diameter. Table 1 has tubing, casing, and gas lift mandrel dimensional data for a clearance check. The 9 5/8” casing string from Figure 3 has a drift diameter of 8.525”, much larger than the major outside diameter (OD) of a 2 7/8” gas lift mandrel (9 5/8” casing is usually paired with 3 1/2” or 4 ½” gas lift completions). The 7” liner has a drift diameter of 5.969” which will clear the 2 7/8” gas lift mandrel major OD of 5.5” for a side pocket or 4.835” for a tubing retrievable option. Downhole data confirms clearance, so attention can be turned to surface facilities.

Table 1: Production Tubing, Casing, and Valve Mandrel Dimensions for Clearance [3,4,5]

Gas compressors, dehydrators, and meters are the crucial complements to the wellbore, subsurface valves, and low pressure gathering/treating facilities of the gas lift system, Figure 4. All operations staff will tell you that gas lift success depends on near 100% run time from compressors, dehydrators, and gas lift gas meters. The rate and pressure available to each well must be adequate and steady. If compressors go down every few days due to poor maintenance or old equipment, then the gas lift system cannot stabilize. Wells are always in a startup (kick off) mode, not steady state operation. If dehydration is malfunctioning or the triethylene glycol (TEG) is so fouled that it cannot absorb water vapor from the gas stream, then hydrate at chokes, regulators, distribution piping, or fuel gas supply lines will cause individual wells, or portions of the field, or the entire field to go offline. Liquid accumulation over years due to condensing water vapor in the piping system can lead to corrosion, liquid slugging, and loss of gas transmission efficiency. Since most gas lift gas is dehydrated but unprocessed solution gas plus returning lift gas, it usually contains heavy hydrocarbon components which can condense and accumulate in the piping system causing problems. 

Figure 4 – Gas lift field schematic with well and surface facility (6)

The critical quality of gas lift gas centers on water content in the gas at compressor discharge pressure and temperature. The value can be estimated from Figure 5. At compressor discharge downstream of the aftercooler, the pressure is 1400 psig and temperature is 120⁰F. The water content point on the chart is 80 lbs water vapor per million standard cubic feet (MMscf) gas. To prevent water condensation down to 40⁰F at 1400 psig, the water content must be reduced to 7 lbs water per MMscf and the TEG dehydrator must remove 73 lb/MMscf. Colder climes often require 1 lb/MMscf to achieve a dew point of ‐10⁰F, which requires 79 lb/MMscf water vapor removal. These estimates set the dehydration requirement to prevent water condensation in the piping which leads to hydrate, corrosion, and water accumulation.

When the dehydrator is out of service for an extended period, or if dehydration is not installed under the false assumption of no problem with hydrate, corrosion, or water accumulation in the injected gas piping system, then a chart is used to predict hydrate formation pressure and temperature based on the gas lift gas specific gravity (ranges from 0.65 to 0.8). Gas lift gas (0.7 specific gravity or relative density) at our example 1400 psig discharge pressure could form a hydrate at approximately 68⁰F, which during normal operation, could occur at pressure expansion (and cooling) locations at chokes, regulator valves, and piping low points where water accumulates. As weather cools many points, including the main pipeline or laterals, could be subject to hydrates. Methanol or ethylene glycol injection stations would be required at potential hydrate points to keep the gas lift operating.

Figure 5 – Water content of hydrocarbon gas (7)
 

Figure 6 – Pressure‐temperature curves for predicting hydrate formation (7)

Section IIIC Gas Lift Well Unloading (Kick off) Guide
An important first step is extracting the control (kill) fluid that is used by the completion/workover team to permit safe installation of the downhole equipment. Even though blowout preventers (BOP) are used, control (kill) fluid in the wellbore tubing and tubing‐casing annulus is the primary barrier to prevent reservoir fluid flow. With the well full of control (kill) fluid, the BOP is removed and replaced with the tree of valves and flanges are sealed. Now thevcritical unloading (kicking off) of the well can be slowly initiated to prevent erosion of valves/mandrels as thevcontrol (kill) fluid passes from the annulus to the tubing, where it flows up to the wellhead to be removed from the well.

Damage prevention to valves and mandrels requires actions prior to installing the downhole equipment and after tree installation. The following practices are applied during the workover to install the packer, tubing, mandrels, and valves:
a) Circulate the wellbore to remove any drilling mud before perforating, running other completion equipment, and installing the gas lift valves.
b) Use a casing scraper to remove debris that adheres to the casing wall and burrs created when packers were set; circulate the casing clean.
c) Use filtered completion and workover fluids and leave filtered fluid in the tubing‐casing annulus. Unfiltered fluids are often a source of solids that can either cut out or plug the gas lift valves.

Unloading the control (kill) fluid from the tubing and annulus is initiated after the well is secured to the production facility:

a) Displace with unloading rates not exceeding 1 barrel per minute (BPM) to prevent erosion of gas lift valves.

b) Start injection gas flow, control rate to attain a 50 psig casing pressure increase in 10 minute increments.
c) Continue this injection rate until the casing pressure reaches 400 psig.
d) Increase the injection gas rate to achieve a 100 psig increase in 10 minute increments.
e) Monitor for an injection gas pressure drop and the return of aerated fluid from the production tubing to indicate gas is injected through the top unloading valve.
f) Observe and record casing pressure (downstream of injection choke or regulator valve) to confirm casing pressure decline as injection point transfers to deeper valves.
g) Use acoustic fluid level tools in the casing annulus to confirm depth of injection.
h) Ensure that injection gas flow is continuous and avoids the occurrence of CRITICAL FLOW where the ratio (P2/P1) of the downstream choke pressure, P2, and the upstream pressure, P1, are in a range well above 0.60., i.e. 0.85 – 0.65.

The depth of injection is related to reservoir pressure (and corresponding tubing pressure) compared to the casing injection pressure. Early operating life may have gas lift injection at a mid‐point in the wellbore, but as reservoir pressure and tubing pressure decline with time, the injection point will automatically shift to a deeper valve where gas injection pressure is greater than tubing pressure. Testing after unloading coupled with nodal analysis simulation from each valve mandrel depth can indicate the point of operation, illustrated in Figure 7. This figure shows the well unloaded to the deep mandrel at 8000’ (Test 1), but a compressor outage caused a shift to the shallow mandrel at 4800’ (Test 2) (the well may not automatically return to the deep point of injection after the restart). Operators adjusted injection rate and another test indicated lift at the 7150’ mandrel (Test 3), and a subsequent test showed a return to the mandrel at 8000’ (Test 4). When the well is unloaded to the depth possible based on available injection pressure, and confirmed with acoustic fluid level and testing, then optimization testing can begin.

Figure 7 – Production rate versus mandrel depth (1)

Section IIID Gas Lift Well Optimization
Optimization based on well tests is an ongoing process since the reservoir inflow performance is continually changing and injection gas needs to be allocated to the best wells in a group supported by the same compressor station. However, confirmation of deep lift (related to available injection pressure) should be done first based on the prior section. The well tests, flowing gradient surveys, and measured flowing bottomhole pressure data from permanent sensors are used to build nodal analysis models that accurately describe the wells’ response to greater or lesser amounts of injection gas. The well performance curve, previously shown in Figure 7, is also called an “optimization” curve and charts gross production (oil and water) rate versus injection gas rate. Each well in the field is tested over a range of 80% to 120% of design injection rate; the results permit choice of operating point based on some criteria: maximum oil, best economic condition, flow stability to minimize slugging, water injection capacity, injection gas capacity. The group of wells in same facility can have injected gas allocated to achieve the “optimum” operating point for each well. Often, optimization is not attained.

When wells falter and optimum points cannot be achieved, troubleshooting techniques are applied to obtain data for problem resolution. Diagnostic techniques, solutions, and problems are addressed in the following table:

Section III Summary
This gas lift tip of the month (TOTM) provides information on gas lift well selection and its inflow relationship with the reservoir, on wellbore and facility parameters that must be addressed for operational success, on the critical stage of unloading control (kill) fluid that is in the wellbore following all interventions to install downhole equipment, and on the optimization procedure plus trouble shooting guides.

Section IIIA links gas lift choice to reservoir gas content since high available reservoir gas requires a lower supplement from gas lift. The effect is density reduction in the wellbore and nodal analysis graphs are used to indicate the difference between a lower natural flow rate and a higher gas lift rate that results from the lower pressure in the bottom of the bore. Nodal analysis with SNAP is used to evaluate interdependence of reservoir productivity, tubing size (with corresponding casing size), and gas lift injection pressure availability.

Section IIIB has an example that compares tubing, gas lift mandrel eccentricity, and casing drift diameters to assure the equipment can installed in the wellbore. Downhole evaluation is necessary, as is a review of surface facilities. Success with gas lift depends on high reliability compressors and dehydrators that provide steady gas capacity without hydrates or liquid accumulation.

Section IIIC provides guides to the crucial step of removing the control (kill) fluid without causing erosional cutting of valves/mandrels. Damage control involves actions prior to installing the wellbore hardware and precautions during the fluid extraction procedure. Displacement of control (kill) fluid is limited to 1 barrel per minute, deemed by testing to be the maximum rate where erosion will not occur.

Section IIID provides the multirate well test optimization procedure to obtain gross production rate versus injection gas rate varied over a range of 80% to 120% of design rate. This test phase begins after the well is confirmed to be lifting at a depth consistent with the available injection pressure. An “optimum” operating point is selected based on criteria such as maximum rate, or economic return, or flow stability, or capacity limits of
injection gas or injection water. When testing reveals a problem, then trouble shooting analysis begins using the guides in this section.

The Authors acknowledge and express our gratitude to Kindra Snow‐McGregor and to Mahmood Moshfeghian for their valuable review and feedback.

REFERENCES

1. PetroSkills Gas Lift (GLI) course manual
2. MEHDI ABBASZADEH SHAHRI, M.S. PhD Texas Tech – 2011
3. Renato Venom Technology and Inspection Services, LLC – 2022
4. Schlumberger Camco Gas Lift Catalog
5. Weatherford Gas Lift Catalog
6. API Gas Lift Handbook
7. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9th
Edition, 3rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

I)  Introduction: Gas – Lift Operational Fundamentals

In the Part 1 of this Series on Gas Lift History and Basic Well Parameters, an attempt was made to bring into focus the primary “state of affairs” of Gas Lift operations in the USA. A discussion was presented related to a candidate Gas Lift well’s completion design that included a typical Casing/Tubing sizing sequence. The function of the production tubing gas lift Mandrels and their function in starting a “kick – off” procedure in a candidate well were discussed. Types of Mandrel Gas Lift Valves were discussed, along with a discussion of the Single Gas Lift Valve employed as the final receptor of the injected casing gas.    

Part 2 will discuss basic Gas Lift well casing and tubing components, and their operational function, as well as Choke Flow relationships in Gas Lift wells.

In the First Section II.A, Energy and Mass Balance relationships will be used to compute flowing pressure gradients, (dP/dL) (psi/ft) for injected casing gas ((dP/dL)g), and for further documents addressing this subject, multiphase flow in the tubing ((dP/dL)mp).

Section II.B will address gas injected at surface into the annular space between production casing and tubing. The injection gas travels down the annular space on its way to either a “kickoff “gas-lift” valve located in a tubing MANDREL with an Injection Pressure Operated gas lift valve (IPO), or to the bottom Orifice GLV. This design was considered and included in the Part 1 discussion. The reader is referred to Part 1 for a description of the IPO. Figure 1 [1] illustrates a representative wellbore configuration for a gas lift well, and the flow paths of the fluids. There are multiple variations of oil well completions for Gas Lift operations that involve various Casing, Liner and Tubing sizes and configurations. As shown by Figure 1 [1], the Gas Lift well “active production” completion consists of an assumed 9 5/8” production casing set a given predetermined depth. There are additional larger casings set at shallower depths, but these are not active in the annular area available for injected casing gas flow. The injected casing gas first encounters the annular space between the 9 5/8” casing and the tubing (assumed to be either 2 3/8”, or 2 7/8” production tubing).  In addition to the 9 5/8” casing, it is assumed that the well is equipped with two liners; a 7 inch liner, and a 5 inch liner. These compliment the assumed 9 5/8” set casing. As shown, the final operating (lowest) GLV is located in the region of the 7” liner. These Well completion configurations may vary from classical Csg/Tbg installations.

Calculations will be performed to determine injected gas annular flow vs. pressure loss related to the 9 5/8” casing and either the 2 3/8”, or 2 7/8” production tubing. The flow is then considered in the annular space between the 7” liner and either the 2 3/8’’, or 2 7/8’’ production tubing. Casing gas flow does not encounter the 5” liner. Physical dimensions for these selections will be addressed. 

Figure 1:  Figure showing vertical section of typical tubing conditions for single phase gas casing/tubing annular flow (1), and two phase  gas/liquid flow.   [1]

In this Section only Gas Flow is assumed, and flowing “Casing-Tubing” gas pressure gradients, (dP/dL)g (psi/ft) will be determined for the selected annular flow configurations.

The calculations will follow the Darcy friction loss correlation within a Bernoulli formulated analysis, along with the Fanning friction factor extracted from the Reynolds Number vs Moody Friction Factor graph (Re vs. f_m), The Fanning friction factor, f_f, is calculated as 25% of the Moody f_m.

Graphs will be presented showing the casing gas flowing gradient, (dP/dL)g (psi/ft) for variations of the injected casing gas rates, (Qg, SCF/D), within the dimensions cited for casing and tubing sizes.  Median flowing conditions will be chosen to represent the flowing temperature, (°F), as well as fluid physical properties for gas gravity, γ_g , and the compressibility factor, Z. An 18.3 lb/lbmol molecular weight gas is assumed for all cases.

Section II.C presents the basic, single phase gas flow performance characterization related to CHOKE FLOW in the Gas Lift Valve. Once a valve has been fitted with a choke (orifice) size, the flow performance of a choke will follow the mass, and energy balance relationships related to isentropic gas expansion.

This flowing condition for the choke MUST be selected in its transitional, sub-critical flow region so that additional changes to injection gas flowrates may be made if called for. It is essential to design the final Orifice GLV so that operation near, or in the Critical Regime (Sonic Flow) is avoided.

As will be shown, the isentropic expansion of an orifice expanding gas is related to a specific pressure volume relationship at constant temperature: PV^k = C, where k is the specific heat ratio. SUBSONIC (transitional choke flow) will be addressed via the Thornhill-Craver [2] Equation (11). This relationship has been shown to match many choke flow measurements. Specific pressures, temperatures and physical properties will be shown.

 

II.A) Basic Mass and Energy Balances.  Gas and Gas-Liquid Flow Pressure Gradients.

The energy balance for gas/liquid flow in either a horizontal or vertical conduit is provided as Equation 1 in its Pressure Gradient form. Figure 1 provides a visual reference for the dimensional criteria for the injected casing gas, or gas/oil flow from the producing Horizon “Pay zone”.                           

      1 [3]

             Flowing Pressure       Friction             Head                 Momentum 

                    Gradient             Gradient          Gradient                Gradient

NOTE: With due consideration, the units selected for the Gas Lift written material will be in FPS units. Conversion factors will be included for assisting the reader. Values will be provided in convenient units to yield appropriate results. 

Where:

dP/dL =   Pressure gradient lbf/ft²ft (psi/ft): (Pa/m)

ρ         =   Density of flowing phase: Single phase gas or two-phase gas/oil – water: lbm/ft³ (kgm/m³)

v         =   Velocity of flowing phase: Single phase gas or two-phase gas/oil – water: ft/sec (m/sec)

d         =   Diameter of production tubing for gas/oil-water flow; Equivalent annular diameter for Casing-tubing annular flow: IDcsg – OD tbg: ft (m)  

Ɵ        =   Well production string angle; (for vertical; Ɵ = 90° (Sine Ɵ = 1))

g_c        =   Gravitational mass – force constant: 32.2 lbm ft/sec²lbf (1 kgm m/sec²N)

g         = Gravitational acceleration, 32.2 ft/sec²

f_f         =   Fanning (Moody) friction factor from Reynolds Number, Re, taken from Moody Friction Factor Chart,

ρ         =   Density of the flowing Fluid: 1) Injected casing gas, 2) Production tubing Oil, Formation Gas (GOR)), and Injection Gas Oil Ratio IGOR, giving the total Gas Liquid Ratio or GLR (SCF/STB):  lbm/ft³ (kg/m³) 

The Reynolds Number, for use with the Moody Friction Factor Chart is defined as follows                               

       2 [3]

Where:  Terms are defined above.

The fluid viscosity, μ, is defined by the slope of the “Shear Stress” vs “Shear Rate” correlation. The units of viscosity are:  1 cp = .001 kg/m sec = 6.72 E-4 lbm/ftsec.

All terms related to the pressure gradient are in an energy base of: lb_f/ft²ft, or the SI equivalent per unit mass (N/m²m). Proper conversions can be easily applied to convert the flowing gradient to typical units of: lbf/in²ft: Pa/m. Notice the first term is the friction gradient, referred to as the Darcy Equation. The “f” parameter is the basic Fanning friction factor that is applicable to the Darcy friction loss term. If selected from the Moody Reynolds Number correlation, the Fanning f_f is the Moody value divided by 4 [3]. The “height” term is represented by the second gradient, which includes the Sine function to consider a non-vertical flow path while the last term is the momentum gradient.

Numerical computations have shown that the momentum component pressure loss does not play a large part in the gradient calculations both for single phase gas, or two phase gas/oil. Thus, the total pressure loss considered is designated by the Friction, and Head Gradients.

In terms of a flowing gradient: (dP/dL); psi/ft (kPa/m), the collected terms are dependent on the density, lbm/ft³ (kgm/m³) of the gas or two-phase gas-liquid, as well as the actual flowing gas, or oil/gas velocity, and “wetted” diameter of the flowing scenario. The total pressure gradient may be determined from these considerations. 

 

II.B)   The case of Casing/Tubing injected gas flow. An effective annular area must be

            determined by:

AREAeff  (in²) =  Casing Inside Area (in²) – Production Tubing Outside Area (in²).

The fluid density term was taken from the standard equation for a gas, applying the following nominal range for the parameters as shown by Figures: A-2 – A-5.

Nominal “Average” Data Range for Vertical Gas Flow in a Gas Lift Well:    

                 T – ºF           100 – 180 °F;       Z  =  .85 – .95

                 P – psia       500 – 2500 psia:   Gas Viscosity – cp    .012 – .018           

                 k (Cp/Cv)   1.23 – 1.24:           MWg = 18.3 lb/lbmol

 Table 1: Nominal parameters chosen for flow equations.

The following data were selected from Ref [4] reporting the commonly used Casing/Tubing dimensions employed in a Gas Lift well completion design:

1.     Data:  ENPRO (OCTG); Casing and Tubing, (OCTG Pipe) [4]

 Table 2: Production Tubing, and Casing design parameters [4]

As discussed, the application of the flowing pressure gradient term requires an effective diameter and Area for tubing and casing. Flowing conditions also involve flow in the annulus between the casing and tubing. An  “effective flow diameter” must be determined from the “effective area” exposed to the gas flow.  Thus the effective annular areas, Aeff, and corresponding effective diameter, Deff, for the Casing–Tubing Configurations would be as follows:

Case 1:  9 5/8 in. Csg / 2 3/8 in. Tubing:  Aeff₁= .785{(8.681)² – (2.375)²} = 54.76 in.²    

Effective Diameter 1: Deff 1 = 8.35 in.

Case 2:  9 5/8 in. Csg / 2 7/8 in. Tubing:  Aeff₁= .785{(8.681)² – (2.875)²} = 52.70 in.² 

Effective Diameter 2: Deff 2 = 8.19 in.

Case 3: 7 in. Csg / 2 3/8 in. Tubing  : Aeff₁=  .785{(6.094)² – (2.375)²} = 24.72 in.²    

Effective Diameter 3: Deff 3 = 5.91 in.      

Case 4: 7 in. Csg / 2 7/8 in. Tubing  : Aeff₂ = .785{(6.094)² – (2.875)²} = 22.66 in.²         

Effective Diameter 4: Deff 4 = 5.37 in.

Table 3: Casing – Tubing Configurations showing Effective Diameter and Area [4] 

The physical dimensions chosen for the well completion will impact the Casing Gas injection rate due to a reduced flow area. This area is specified as an “Effective Area”, “Aeff” in.². This reduced area will impact maximum flow in the casing/tubing annulus.

Excessive gas flow velocities are not recommended, as friction gradient increases rapidly, and impact optimal oil flow. As shown, flow from the surface to the final GLV would generally be a configuration of the above selections, depending on gas rates, and depth. All these conclusions are essentially based on the “assumptions” stated for the Casing / Tubing dimensions.

Equation 3 [3] indicates the “head” or hydrostatic term for a gas column, expressed in pressure terms, is given by:

                                               ΔP_gas = ρ_g(g/gc)ΔH                                  3 [3]

When integrated over a differential height, the static gas pressure at the bottom of a vertical column Can be calculated as:              

                                                   P₂ = P₁e^s                                                       4 [3]

Where :

P₂      =     Pressure at bottom of column  – lbf/ft²

P₁     =     Pressure at top of column – lbf/ft²       

e       =     Naperian Log base : 2.718

s       =     Correlating function :  ΔHγg/AT_mz_m  

ΔH   =     Height differential – ft

γ_g     =      Gas gravity          

T_m   =      mean absolute temperature – °R

Z_m   =      mean value for z.       

Fortunately, the generalized equation for computing the total flowing gas gradient in a conduit at “median” conditions for all fluid dependent terms, including the vertical head term are available. We have defined a simple and direct approach via:

                                                               5  [3]

To further simplify the actual calculations, an assumption has been made to determine the Moody friction factor by a simplified equation for f, assuming turbulent flow. This equation is based on small pipes [3]. It has been sparingly used in gas flow “f” factors; however in most Gas Lift cases, the injected casing gas is in the turbulent flow range, and resulting f_{MOODY ,} evaluated for small conduits (3” – 8”) is assumed to be applicable: For  application  of  the   flowing    gradient equation, the Fanning friction factor f_f  is applied. The Fanning friction, f_f,  factor is the f_MOODY divided by 4.                               

                                   f_MOODY = 4 x 0.0239(Re^{– 0.134})                                    6  [3]

The calculation procedure will be based on the follow the sequences:

 II B.1) Calculation of flowing Gas gradients:

Curves are presented in Appendix A showing the flowing casing gas annular flow pressure gradient, (dP/dH)csg in psi/ft for the following sequence of data :

Series 1:  9 5/8”  Casing with 2 3/8” production tubing

Series 2:  9 5/8”  Casing with 2 7/8” production tubing

Series 3:  7 ”       Liner with   2 3/8” production tubing

Series 4:  7 “       Liner with   2 7/8” production tubing

The gradient data will be simplified for the casing/tubing gas flow indicated by the following input flowing physical properties, and conditions:

• Surface Injection Pressure = 500, 1000, 1500, 2000, 2500  psi
• Flowing temperature range = 100 – 180 °F
• Mean “ z” factor  = 0.88
• Molecular Wt. Gas = 18.3  lb/lbmol
• Mean Gas Viscosity =  .014 cp (9.4 E-6 lbm/ftsec) ; ( 1.4 E-5 kgm/msec)
• Aeff, Deff = Effective Casing-Tubing Flow Areas (in²), and Diameter, Deff, (in.) for annular flow yielding an equivalent effective “flow area” are taken from Table 3.
• Selected Flow Range : 0.5 – 2.5 MMSCFD

As can be surmised, the gradient flowing parameters for pressure “gain” in the existing Casing/Tubing annulus must be known as this information yields the existing pressure at depth of the final (lowest) orifice Gas Lift Valve, shown in Figure 2.

Figure 2: Bottom Hole Production Tubing Gas Lift Valve [6]

Notice from the presented data, that a reasonable pressure gradient range is duly established by Figures [A.1] through [A.4], Appendix A. As shown by the Figures, the flowing gradients range discreetly between 0.01 and 0.06 psi/ft.  For a set casing/liner size, (either    9 5/8” or 7”), the production tubing diameter, i.e. from 2 3/8” to 2 7/8” will impact flowing gradients. The shown “Effective Area, and Diameter” are seen to decrease discretely, showing a slight increase in the flowing gradient. It is interesting, and important to note that beyond a casing injection rate of 0.5 MMSCF/D, the flowing gradient shows to be essentially constant. This occurrence is due to the impact of the flowing gas density term which is included in the friction loss portion of the gradient with a negative sign, and also is part of the corresponding increase of the static gradient which is positive. The curves should be very applicable when estimating the Casing gas injection pressure as illustrated by the following example:

Example 1: Gas Lift Well –  Casing / Production Tubing Gas Injection:

a)     Depth of pay zone in candidate well:  8000 ft.

b)    Csg/Tbg upper completion: 9 5/8” x 2/3/8”: Depth: 7000 ft. 

c)     Liner/Tbg. Lower completion:    7” x 2 3/8”: Depth:  8000 ft.

d)    Injection Gas Rate :  1.5 MMSC/D

e)     Well Head Injection Pressure, and Temperature: 1500 psia, :  140 °F

f)     Mean Z factor : 0.85

g)    Mean gas viscosity:  0.015 cp. 

h)    Mol. Wt. Gas : 18.3 lb/lbmol 

1)    Solution : 

a)    Gradient in 9 5/8” x 2/3/8”annulus : 0.035 psi/ft  (Figure A-1)

b)    Flowing Pressure at 7000 ft.  : 1500 + (0.035)(7000)  = 1745 psia

c)     Gradient in 7”x 2 3/8” annulus: 0.038 psi/ft   (Figure A-3)

d)    Production Tubing Valve Injection Pressure: = 1745 +(0.038)(1000) = 1783 psia 

With these pressure vs depth data, and a knowledge of the well’s formation IPR, as well as the Solution Gas Oil Ratio (GOR – SCF/STB) of the formation fluid, the conditions are now set for a selection of the proper ORIFICE SIZE to be installed in the Production Tubing GLV.

II C)  Consideration of Basic Orifice Flow as applied to an operational Gas Lift Valve (GLV)

To consider the fundamental conditions governing “orifice” flow, the basic energy balance is once again addressed. However, for the orifice flow case, the equation does not include the “friction loss” term, or the “head” term. The equation now considers the “pressure change” solely  as a function of the momentum term:

                                                                      7   [3]

For the specific case of an assumed adiabatic reversible (isentropic) expansion, pressure and specific volume V, (ft³/lbm), and pressure (lbf/ft²) are related by:

                                          PV^k        =    Constant                                                    8 [3]

In Equation 8, the exponent, k, is the ratio of specific heats, Cp/Cv. With this assumption, the pressure loss across a restriction “choke”, could be fully recovered and downstream pressure would equal upstream pressure.  In reality, some energy is lost to friction and the process is not reversible.  Typically, this non-ideal behavior is resolved by assuming ideal behavior, then applying a correction factor to the result, which is the approach we will use here.  The Thornhill – Craver relationship, Equation 11, applies this concept with the inclusion of the Cd factor. Alternatively, the exponent can be modified to better represent reality (such that “k” is not equal to Cp/Cv).  In the latter case, the exponent is often determined experimentally, and is often referred to as the “polytropic” exponent.  Either approach allows us to consider the case where the recovery of the upstream pressure is not reached due to some “polytropic” effect.  The magnitude of the deviation is very dependent on the precise physical configuration of choke internals.

Notice in Figure 3, [5] depicting an orifice (choke) in a flowing condition, that the exchange of energy results in an increased “velocity”, and reduction of pressure for the fluid as it passes through the choke (orifice). The point of minimum pressure is termed the “Vena Contracta”. When the downstream area increases, the flow velocity decreases, and flowing pressure recovers; however due to the polytropic (i.e. non-ideal) conditions the recovering pressure does not reach the inlet value. This is the polytropic “energy loss”.

Figure 3: Flow through an orifice (choke) for ploytropic conditions  [5]

Equation 9 now becomes:

                                                                              [3]

This is the basis for most choke performance analyses, relationships, and working equations. This relationship yields the equations applied by Gilbert [6], Fortunati [7], and others. The fluid velocity, v, (ft/sec) can be related mathematically to the pressure differential, or pressure ratio across the choke.

The importance of all these solutions is the fact that a plot of downstream to upstream pressure ratios for choke flow: P₂/P₁ yields the application for fluid flow deemed the “transient, subcritical” flow regime. As shown by Figure 4 [8], for a given upstream pressure, P₁, the flow increases across the orifice as downstream pressure, P₂ is reduced, until further reduction of P₂ will no longer increase flow. The reason for this behavior is that pressure disturbances are transmitted “upstream” at the speed of sound. Thus, when this velocity of the fluid at the vena contracts reaches the speed of sound, no additional flow is possible.  This is referred to as CRITICAL flow. In this case, the “critical pressure ratio” is reported as 0.588, however this value is not a constant for typical natural gas flow through a choke. The critical pressure ratio varies from approximately 0.49 to 0.59.

Figure 4: (P₂/P₁):

in an Orifice, or Choke [6]
                       

For the purpose of this document, the Thornhill-Craver [2] Choke flow relationship has been applied. This relationship addresses the non-ideal (non-reversible) behavior in part by applying a correction factor, in this case, C_D, the discharge coefficient which is derived experimentally.

In many Gas Lift instances, the calculated “choke capacity” has been well matched by this relationship. Equation 11 is the standard Thornhill-Carver giving a choke (orifice) capacity in SCF/d with all the parameters involved in the polytropic energy flow relationship. Appendix B presents the choke capacity results for various cases that coincide with the Flowing gradient curves with Appendix A. The selected data ranges are:

Upstream Pressures : 1000, 1500, 2000, 2500 psia

Choke Diameters : 0.125” (8/64ths”), 0.25” (16/64ths”), .375” (24/64ths”), .50” (32/64ths”)

Inlet Temperatures:  150, 175, 200, 225  °F

Gas Gravity : 0.63 (constant) – MW = 18.3 lb/lbmol

Z = .87 – .88 -.89  :     k = 1.235 , 124 , 1.25 , 1.26 

Equation 11, the Thornhill – Craver correlation will compute the specific choke capacity as a function of Pressure Ratios related to choke flow: P₂/P₁, downstream / upstream.  The Figures in Appendix B show the flowing gas capacity as a function of the pressure ratio. All cases are seen to show the decrease in pressure ratio with increase in flow UNTIL the critical point is reached.  This value is in the range of 0.55 – 0.48. Thus, once a bottom hole injection pressure has been determined, the choke must be selected to coincide in capacity (SCF/D) within its acceptable “sub critical” flow regime. 

Equation 11 [2] is the Thornhill-Craver for Choke Flow through an orifice.

  11 [2]
 

Where:

Qg=gas-flow rate at standard conditions (14.7 psia and 60°F), Mscf/D,Cd=discharge coefficient (determined experimentally), dimensionless,A=area of orifice or choke open to gas flow, in.2,P1=gas pressure upstream of an orifice or choke, psia,P2=gas pressure downstream of an orifice or choke, psia,g=acceleration because of gravity, 32.2 ft/sec2,k=ratio of specific heats (Cp/Cv), dimensionless,T1=upstream gas temperature, °R,Fdu=pressure ratio, P2/P1, consistent absolute units The correlations presented will allow the reader, and designer of a Gas-Lift well/system to choose a proper choke based on the flowing predictions for 8, 16, 24, and 32/64ths in chokes. The selected data will be displayed on the graphs. Rarely have choke performance curves been presented in this fashion, and their applicability is proven in that at the critical pressure ratio, at approaching sonic velocity, can be shown to exist. This critical pressure ratio has been shown to be expressed by: 12  [2] Where : Fcf   = Critical Orifice pressure ratio: P2(vena contracta) / P1 (inlet) k      = Specific heat ratio: Cp / Cv All assumed parameters have been chosen to represent actual Gas Lift wells in the field. Further, the parameters are extended so the impact of critical flow can be seen.to.If the previous example is revisited: Example 2: Gas Lift Well –  Casing / Production Tubing Choke Capacity: a.     Depth of pay zone in candidate well:  8000 ft. b.     Csg/Tbg upper completion: 9 5/8” x 2/3/8”: Depth: 7000 ft.  c.     Liner/Tbg. lower completion:    7” x 2 3/8”: Depth:  8000 ft.d.     Injection Gas Rate :  1.5 MMSC/De.     Casing Injection Pressure  = 1500 psia Solution: Flowing pressure at Production Tubing/ 7” Liner depth is found to be: 1873 psia Referring to Appendix B, two (2) Choke Dimensions are seen to provide the desired solution for the correct flowing bottom hole liner/tubing capacity in the orifice valves shown as follows: As shown from the solution, both a 16/64 ths, or 24/64ths Choke will perform adequately. Perhaps the larger size, 24/64ths is the correct solution, as the remaining subcritical flow regime is greater: 0.96 – 0.5. This will allow for more casing gas to be injected, and provides leeway for corresponding increases, or reductions to the choke inlet pressure as flow conditions develop.  Summary: In this Tip of The Month an analysis was performed in Section II A to review energy and mass balances as related to a candidate Gas Lift Well’s flowing gradients. Standard Energy balance equations provided the proper data to simulate both the annular flow in a casing/ tubing configuration, as well as for choke performance where energy change was related to momentum differential. In Section II B, Casing/Liner, and Tubing sizes were selected to perform the necessary flowing gradients for the injected casing gas.  Equipment Data recompiled from Industry standards provided the necessary injected Casing/Tubing/Liner dimensions to calculate Effective Areas, as well as Effective Diameters for flow of the injected casing gas to the Production Tubing Gas Lift Valve at a given depth. Section II C reviewed the existing Industry standards for Gas Lift Choke flow. The Thornhill–Carver equation provided the choke performance data relating to the selected choke sizes. An example was presented to provide a flowing liner/tubing bottom hole pressure from the Gradient curves presented in Appendix A. The Choke performance curves in Appendix B were employed to “size” the proper choke diameter, and flowing conditions. The Appendix B curves apply to orifice sizing but are only an estimate for gas lift valves since the valve stem in the seat reduces flow area. To learn more about similar cases and how to minimize operational problems, we suggest attending the G-4 Gas Conditioning and Processing Course presented worldwide. In addition, our session in ALS (Artificial Lift Systems), GLI, Gas Lift, and Completion and Workover courses.             By:  Frank E. Ashford, Ph.D. Wes H. Wright The Authors would like to cite and appreciate the support received from Mr. John Martinez, PetroSkills Senior Advisor. Mr. Martinez is a known expert in Gas Lift applications. His excellent technical expertise, and proven field knowledge are reflected throughout this document. We also acknowledge and express our gratitude to Kindra Snow-McGregor for her valuable review and feedback. References: 1.     Simplified and Rapid Method For Determining Flow Characteristics of Gas Lift Valves (GLV); Mehdi Abbaszadeh Shahri: Texas Tech University, Aug. 2011 2.     Cook, H. L. and Dotterweich, F.H. 1946 Report on Calibration of Positive Flow Beans manufactured by the Thornhill – Craver Company, Houston, Tx. 3.     Gas Conditioning and Processing: The Basic Principles: (John M. Campbell): Vol 1 9th Edition. Editors R.A. Hubbard, K.S. McGregor 4.     ENPRO (OCTG); Casing and Tubing, (OCTG Pipe) : (data and dimensions) 5.     NEUTRIUM: Calculations of Flow through Nozzles and Orifices ; Feb. 11 , 2015 6.     ELSEVIER: Flow Measurement and Instrumentation: Vol 76, Dec. 2020 – David A. Wood , et. Al. Appendix A : Casing – Production Tubing Flowing Gradients  Figure A.1: Injected Gas Flowing Gradients in 9 5/8” Casing and 2 3/8” Production Tubing   Figure A.2: Injected Gas Flowing Gradients in 9 5/8” Casing and 2 7/8” Production TubingFigure A.3: Injected Gas Flowing Gradients in 7 ” Liner and 2 3/8” Production TubingFigure A.4: Injected Gas Flowing Gradients in 7 ” Liner and 2 7/8” Production TubingFigure B.1: Gas Lift Choke Performance: .125 in.  ( 8/64”) Choke DiameterFigure B.2: Gas Lift Choke Performance: .25 in.  ( 16/64”) Choke DiameterFigure B.3: Gas Lift Choke Performance: .375 in.  ( 24/64”) Choke DiameterFigure B.4: Gas Lift Choke Performance: .50 in.  ( 32/64”) Choke Diameter

Hydrogen sulfide and carbon dioxide are the principal objectionable acid gas components often present in natural gas, synthetic gas, and various refinery gas streams. These acid gas components must be removed for corrosion prevention in gas pipelines, process equipment, and for health and safety reasons. Reference [1] provides current acceptable concentration levels for these acid gases in various gas streams. Hydrogen sulfide removal often requires the production of sulfur in the sulfur recovery units to meet emission limits.  Sulfur is a product used to produce sulfuric acid and fertilizers. Carbon dioxide removal is used for enhanced oil recovery and is required for carbon capture and sequestering (CCS) operations.

In natural gas treating, there are several processes available for removing the acid gases. Aqueous solutions of alkanolamines are the most widely used [1]. The alkanolamine (amines process is characterized as “mass transfer enhanced by chemical reactions” in which acid gases react directly or react through an acid-base buffer mechanism with the amine to form a nonvolatile ionic species. For further details of sour gas treating, please refer to references [1-5].

Several amines have been used for acid gas removal from natural gas streams. In this study only a tertiary methydeithanolamine (MDEA) will be considered. MDEA has the highest selectivity and lowest regeneration duty required, making it the solvent of choice for many applications. Due to the complexity of acid gas removal process and the corrosivity of the rich amine solutions, there are many operating problems that can occur.  Likely the most common operating problem in these systems is amine foaming, which often results in failure to sweeten the inlet sour gas to the required specifications. The details of multiple operating issues and how to prevent them are discussed in PetroSkills-Campbell’s Gas Conditioning and Processing, G4 and G6 Courses (Gas Treating and Sulfur Recovery), as well as in our new Acid Gas Removal – Amine Focus Fundamentals three-day short course.

Foaming is caused by several factors, but the most common in gas processing plants are the condensation of liquid hydrocarbons, ingress of other contaminants in the process gas stream to the amine contactor, and amine degradation products. Symptoms of foaming include a sharp rise in contactor differential pressure, reduced rich amine flow leading to loss of liquid level in the contactor bottoms and severe carryover of amine into the outlet scrubber. It is nearly assured that the treated gas will go off-spec during a foaming event. In this TOTM, we will study the failure to sweeten due to the presence of hydrocarbon liquids in the feed sour gas. It is important to remove the hydrocarbon liquids from the inlet gas stream to prevent solvent contamination. In addition, any hydrocarbon liquid condensation must be avoided.  The inlet gas hydrocarbon dewpoint increases as a result of the acid gas components being removed from the gas stream in the amine contactor.  Refer to Chapter 4 of Reference [2] for more information. Thus, it is important to understand the hydrocarbon dewpoint of the gas stream as it flows up the contactor to ensure that the amine feed temperature is above the treated gas hydrocarbon dewpoint condition to prevent condensing liquid hydrocarbons in the amine. Predicting the hydrocarbon dewpoint of the treated gas stream requires a thorough understanding of the gas composition, including accurate heavy end characterization.

Case Study:

For illustration, we will consider sweetening of 1.415 x 10⁶ std m³/d (49.95 MMSCFD) of a sour natural gas with the composition, pressure, and temperature presented in Table 1. ProMax [6] simulation software with “Amine Sweetening – PR” property package was used to perform all the calculations. In addition, the mass + heat transfer option was used.

Table 1. Feed composition, volumetric flow rate and conditions

The following specifications/assumptions were made:

Contactor Column

a. Feed sour gas is saturated with water

b. Number of actual stages = 23

c. Pressure drop = 35 kPa (5 psi)

d. Lean amine solution temperature = Sour gas feed temperature + 5.56 ℃  (10 ℉)

Regenerator Column

a. Number of actual stages = 22 (excluding condenser and reboiler)

b. Rich solution feed temperature = 98.9 ℃  (210 ℉ )

c. Rich solution feed pressure = 105 kPag (15.2 psig)

d. Condenser temperature = 48.9 ℃  (120 ℉ )

e. Pressure drop = 35 kPa (5 psi)

f. Bottom pressure = 138 kPa (20 psig)

g. Reboiler duty = Specified “Steam Ratio” times circulation rate (Refer to Table 2)

Heat Exchangers

a. Lean amine cooler pressure drop = 21 kPa (3 psi)

b. Rich side pressure drop= 41 kPa (6 psi)

c. Lean side pressure drop = 35 kPa (5 psi)

HP Pump

a. Discharge Pressure = Sour gas feed pressure + 35 kPa (5 psi) + ΔP_Static

ΔP_Static = At least 130 kPa (19 psi) in just the elevation difference between the pump and the contactor inlet, plus some pressure drop on the FCV (unless you control flow via pump speed).

b. Efficiency = 65 %

Booster Pump

a. Discharge Pressure = Suction pressure + 55 kPa (8 psi) pressure drops in the L/R exchanger and lean amine cooler and ensures the HP pumps have enough NPSHA.

b. Efficiency = 65 %

Lean Amine Circulation Rate and Concentration

a. Varied to meet the target acid gas loading in the rich solution shown in Table 2

Rich Solution Expansion Valves

a. ΔP in the expansion valve to rich amine flash drum (valve 100) = 6310 kPa (915 psi)

b. ΔP in the expansion valve to the stripper (valve 101) = 440 kPa (64 psi)

A simplified process flow diagram for the case studied is presented in Figure 1 [6].

Table 2. Specified amine concentration, target rich solution acid gas loading, and steam ratios [4]

Figure 1. Simplified process flow diagram for an amine sweetening unit [6]

Results and Discussions:

Based on the description and specifications presented in the previous section, the process flow diagram in Figure 1 was simulated using ProMax BRE Software [6]. The simulation was performed for the steam ratio presented in Table 2. The results of the simulation are listed below:

– Treated sweet gas: H₂S (3.98 ppmv) and CO₂ (0.948 mol %)

– Lean amine circulation rate: 669 std gpm (152 std m³/h)

– Reboiler duty: 35.1 MMBtu/hr (10.28 MW)

– Amine make-up: 0.315 lb/hr (0.143 kg/h)

– Water make-up: 646 lb/hr (293 kg/h)

The amine make-up is required to compensate for the amine vaporization losses from the top of the contactor, regenerator, and rich amine flash drum.  The Promax program shows a blowdown stream on the amine, however it should be noted that this is not typically done in the field.  Amine reclaimers are utilized today to recover amine solution quality rather than bleed and feed approach that was used in the past.  The water make-up is required to maintain the water balance as the acid gas and flash gas remove water from the system.  The operating conditions in the contactor will result in condensation or vaporization of water, but this is of really no concern as the amine solution is aqueous. Overall, water is generally lost from the amine solution in the acid gas stream and makeup water is required. This makeup water must be free of minerals and other contaminants, or it may become a means of introducing agents that cause foaming.

Figure 2 shows the feed sour gas phase envelope on dry basis. The calculated cricondentherm, cricondenbar pressures and temperatures are shown in Table 3.

Figure 2. Feed sour gas phase envelope.

Table 3. Sour gas calculated cricondentherm and cricondenbar pressures and temperatures

The gas from the facility inlet separator flowing to the AGR unit is at its hydrocarbon dewpoint at 90°F (32.2 °C) at pressure of 1000 psig (6,897 kPag). Any cooling occurring in the piping between the inlet separator and the process will result in hydrocarbon liquid and water condensation. Because the feed gas pressure is above the cricondentherm pressure, additional liquid hydrocarbons may be formed due to retrograde condensation associated with pressure drop in the piping. In addition, the gas leaving the inlet separator may also contain a small amount of liquid hydrocarbons and produced water due to separator carryover.

Installation of a properly sized filter coalescer upstream of the acid gas removal unit (AGR) unit is recommended, as shown in Figure 3 for an amine unit to minimize any potential liquid hydrocarbon or produced water carryover from entering the amine system. In some facilities the filter coalescer may be installed at a considerable distance from the contactor. When this happens, the condensation illustrated in Figure 2 can lead to solvent contamination and significant operating problems.

Figure 3. Process Flow Diagram for an Amine System Inlet [2]

To prevent hydrocarbon liquids condensation in the amine contactor, the lean amine inlet temperature should be roughly 5 to 10 °F (3 to 5 °C) warmer than the inlet sour gas. This temperature difference is referred to as the approach temperature (lean amine temperature approach to the inlet sour gas).  Figure 4 presents the twenty-three actual-stages temperature profiles in the contactor column for two approach temperatures of 5, and 10 °F (2.8, and 5.6 °C)

Figure 4. Contactor gas-phase stage temperature profiles for two lean amine approaches to feed sour gas temperature

To check for possible hydrocarbon liquids condensation on a stage (or actual tray), the phase envelopes for the feed sour gas and the gas streams (on dry basis) leaving a few selected stages in the contactor column are illustrated in Figures 5 (FPS) and 6 (SI).

Figure 5-FPS. Process gas phase envelopes on dry basis in absorber (contactor) column for feed and selected stages.

Figure 6-SI. Process gas phase envelopes on dry basis in absorber (contactor) column for feed and selected stages

Note that as the acid gases are removed from process gas during its upward flow through the amine contactor, the gas hydrocarbon dewpoint curve shifts to the right and the two-phase region is expanded.  Determining the required lean amine temperature approach requires knowledge of the phase envelopes of the gas as it flows through the contactor.  Maintaining an adequate lean amine inlet temperature above the hydrocarbon dewpoint of the treated gas stream can aid in preventing foaming due to condensing liquid hydrocarbons.

Table 4 presents the estimated temperature values for a few selected points of Figures 5 and 6 based on the specified lean amine approach temperatures of 10 and 5 °F (5.6 and 2.8 °C) at the top of contactor and the feed sour gas temperature of 90 °F (32.2 °C).

Table 4. Estimated temperature margins for two lean amine approach temperatures

Notice that in the 10 °F (5.6 °C) case the stage operating temperature is to the right of (warmer than) the hydrocarbon dewpoint temperature by 8.0 °F (4.5 °C), thus no hydrocarbons vapor from gas stream condenses and solvent contaminations do not happen.

This is not the result for the 5 °F (2.8 °C) case (Figure 7).  The inlet amine solution is slightly warmer than the hydrocarbon dewpoint temperature of the treated gas by 3.0 °F (1.7 °C) with a low safety factor. The potential would exist for some liquid hydrocarbon condensation if the amine were to cool slightly or if additional acid gases were removed from the gas stream.

Figure 7. Gas-phase stage temperature profile and process gas phase envelopes on dry basis in absorber (contactor) column for feed and selected stages

In a number of facilities, the expansion of the treated gas phase envelope and inadequate lean amine inlet temperature (too cool) was the primary cause of liquid hydrocarbons condensing in the amine solvent resulting in severe operating problems.  This will, result in foaming and potential failure to sweeten the gas adequately.  Notice in the case with the smaller temperature approach, one would have assumed that the amine solvent temperature would be sufficiently high enough to prevent liquid hydrocarbons from condensing out.  In this example, the 5°F (2.8°C) approach temperature on the lean amine had a relatively small safety margin above the hydrocarbon dewpoint of the gas stream.  If the amine temperature was not tightly controlled and it entered the amine contactor at a temperature 3°F (1.7°C) cooler, then hydrocarbon liquid condensation would likely occur.   Without detailed knowledge of the treated gas phase envelope and doing this type of analysis one could easily assume that the liquid hydrocarbons were entering the system from the inlet filter coalescer upstream of the AGR.

Conclusions:

Based on the results obtained for the case study considered in this TOTM, we have arrived at the conclusions below.

To prevent solvent contaminations and foaming the following suggestions can be utilized:

a. Install a properly sized scrubber followed by an adequately sized filter coalescer directly upstream of the amine contactor column.

b. Select an appropriate approach temperature to feed gas temperature for the lean amine at top of contactor based on the treated gas composition phase envelope.

c. Plot the process gas streams hydrocarbon dewpoint curves for selected stages and make sure they are located to the left of operating temperatures (Figures 5 and 6).

d. To avoid hydrocarbon liquid in contactor, select the lean amine temperature to yield ΔT_min= or > 5.4 °F (3 °C); where ΔT= Stage gas phase temperature – gas phase cricondentherm temperature is the minimum for stage 1 (sweet gas) or ΔT_min = ΔT_stage1.

To learn more about similar cases and how to minimize operational problems, we suggest attending our G6 (Gas Treating and Sulfur Recovery), G4 (Gas Conditioning and Processing), PF81 (CO₂ Surface Facilities), PF4 (Oil Production and Processing Facilities), courses.

References:

Maddox, R.N., and Morgan, D.J., Gas Conditioning and Processing, Volume 4: Gas treating and sulfur Recovery, Campbell Petroleum Series, Norman, Oklahoma, 1998.

Campbell, J.M., Gas Conditioning and Processing, Volume 1: The Basic Principles, 9^th Edition, 1^st Printing, Editors Hubbard, R., and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.

Campbell, J.M., Gas Conditioning and Processing, Volume 2: The Equipment Modules, 9^th Edition, 1^st Printing, Editors Hubbard, R., and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, 2014.

GPSA Engineering Data Book, Section 21, Volume 2, 13^th Edition, Gas Processors and Suppliers Association, Tulsa, Oklahoma, 2012.

Moshfeghian, M., Bell, K.J., Maddox, “Reaction Equilibria for Acid Gas Systems, Proceedings of Lawrence Reid Gas Conditioning Conference, Norman, Oklahoma, 1977

ProMax 5.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2021

1.0         Introduction

LNG and gas processing plants often have some units that are identical; viz acid gas removal unit (AGRU), mol sieve dehydration unit (MSDU) and mercury removal unit (MRU). The degree of gas sweetening (AGRU) and the degree of water dew point (MSDU) is dependent on the requirement i.e., is the gas going to be transported in pipeline or is the gas is going to be further processed for LNG production. If mercury is present in the inlet gas, mercury removal would be mandatory due to both health and environmental reasons and aluminium metallurgical aspects.

In this TOTM, a closer look is taken on the possible causes of failure or of operational issues of the three processes mentioned above. Further, possible solutions are given in addressing these operational issues. But before going any further, let us consider what needs to be discussed prior to design, engineering and construction of a gas processing facility, or an LNG plant.

 

 

2.0         Some Typical Issues at the Design Stage of a Project [1]

At the design stage of a project, a critical review needs to be taken on the following aspects:

1. Has the upstream feed gas processing been adequately defined?

2. Have the contaminants/impurities levels been properly analyzed, discussed, and documented?

3. Is there sufficient understanding of the feed gas composition or processing steps?

4. What are the overall plant products slate and their specifications? In other words, is the economic case adequately addressed?

5. Have all environmental constraints been adequately identified and addressed in the design?

During this project stage, do not consider this as a stand-alone concept (think: plant life to be 25 years!). Some interfaces with upstream and downstream facilities needs to be defined and it is important that unknown parameters attributed to start-up, end-of-life, and are identified/discussed, and their potential impacts assessed.

 

 

2.1         Some Typical Concerns and Risks at the Design Stage

Having worked on many gas projects, the author has identified the following crucial pitfalls at the design stage of a project.

Lack of understanding/knowledge of upstream processing units at summer/winter periods and normal/upset conditions. These could have significant impact on the overall design. Typical problems to address are [1]:

– Liquid slugs, formation water (think: what water cut is expected over 10, 20 years).

– If a multi-phase pipeline is used to handle gas, liquid hydrocarbons, and water, then handling of large quantities of liquids during pigging operations could impact operations of inlet receiving facilities and consequently handling of condensate handling facilities such as the 3-phase separator and condensate stripper (in some locations referred to as the condensate stabilizer).

– Undersized separators, particularly 3-phase separators at the inlet receiving facilities are a common problem.

– Presence of chemicals (glycol, methanol, corrosion inhibitors, hydrate inhibitors, compressor lube oils, etc) due to upstream operations such as field dehydration, etc.)

– Knowledge of sour feed gas heavy hydrocarbons distribution in the feed gas

– Feed gas composition needs to be analysed with complete breakdown to C12 to C20 fractions. It is not sufficient to have C5+ lumped together as one component for sizing the inlet equipment and liquid handling facilities for hydrocarbon rich gases.

– Presence of aromatics should be analyzed, as these components have a significant impact to the inlet fluids phase behavior and influence the design of amine regenerator and lean/rich amine heat exchanger.

– Inadequate or poor knowledge of sulphur compounds “traces” or distribution e.g., COS, mercaptan species, CS₂, etc.  If the gas is sour, these components should be analyzed to determine their presence and quantities as these can impact the selection of the AGRU.

– Gas composition ranges not identified, for example the H₂S/CO₂ max and min ratios; this dictates the Sulphur Recovery Unit line-up and can have a significant impact on the capital investment costs.

It is the responsibility of the Process Engineer to highlight the critical items that can impact the design and to propose reasonable design limits

 

 

2.3         Block Diagrams of Gas and LNG Plants

The following two figures show the block diagram of a typical gas plant and of a typical LNG plant.

 

Figure 1: Typical Process Block Diagram of a Cryogenic Gas Processing Facility

 

 

Figure 2: Typical Process Block Diagram of an LNG Plant

 

 

In the case of Figures 1 and 2, the commonality of the three basic processing units: AGRU, MSDU and MRU are illustrated. Failure in any one of these units would hamper the operations of any downstream process units and if the problem in any one of the units is not rectified, plant capacity would need to be decreased or a total plant shutdown initiated.

 

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COMMON OPERATING PROBLEMS

 

3.1         Troubleshooting AGRU [2,3]

 

Figure 3: Example Process Flow Diagram for an Acid Gas Removal Unit [5]

 

 

There are three most common classes of problems in AGRU:

– Causes of failure to sweeten: H₂S, CO₂ and total sulphur treated gas specifications are not being met. This could be due to inadequate regeneration of amine due to too high a circulation rate, insufficient reboiler heat input, reboiler problems, plugged or damaged regenerator trays or a leaking lean/rich amine exchanger.

– Foaming: Problem most often occurs at the contactor. Symptoms of foaming include a sharp rise in contactor differential pressure, loss of level in the contactor bottoms and severe carryover of amine into the outlet scrubber. It is almost certain that the treated gas will go sour whenever foaming occurs.Foaming is caused by a number of factors, but the most common in gas processing plants is thecondensation of liquid hydrocarbons or ingress of other contaminants in the process gas stream to the amine contactor.

– Amine losses: The cost of amine can be a large operational expense, especially when it seems to disappear continually. Amine loss is an operational fact of life. When amine losses become excessive, however, it is a problem that requires troubleshooting. While normal amine loss is different for each of the different amines, losses in all systems occur due to the following four reasons [3]:

1. Mechanical loss: means that solution is physically lost from the system either due to carryover, foaming, and simple leaks at fittings and pump seals. Such losses are minimized with mist eliminators, foam prevention and the existence of a separate amine drain and recovery system

2. Solution Conditioning Loss:  Practically a sub-category of mechanical loss, conditioning losses occur when a filter or reclaimer is not properly drained of solution during maintenance or change out activities. This type of loss is minimized by following proper procedures for solution displacement before draining and opening this type of equipment.

3. Chemical and Thermal Degradation: occurs when amines come into contact with oxygen and other components in the produced sour gas. Improper storage or oxygen in-leak are the main causes of oxygen contamination. Some amines are more susceptible than others to reactions with COS, CS₂ and other sulphur compounds. All amines can degrade thermally if subjected to excessively hot heat transfer surfaces. Thermal degradation can be substantial at temperatures above 150°C (300°F).

4. Vaporization Loss: This type of loss depends upon amine type. For instance, MEA is 30 times more volatile than DEA, so we can expect vaporization losses in MEA systems to be substantially greater. Vaporization losses are usually minor and occur at both the contactor and the regenerator of the AGRU.

 

3.1.1     Foaming in AGRU [3]

There are innumerous papers and technical publications on foaming. Here a brief overview of potential foaming causes are given:

– Dirty amine containing iron sulphide or diatomaceous earth.  Both are foam promoters, as are well treatment chemicals that find their way into the system due to insufficient upstream separation.

–  If corrosion inhibitors are in use, they may be the culprits as these inhibitors are surface active species. One of the most common promoters of foaming can be (surprisingly) anti-foam products that are injected to stop foaming.

– Foaming is predominantly prevented by all those things already discussed in other sections, like preventing hydrocarbon condensation, using a filter/separator on the inlet and keeping the solution clean.

– Once it starts, however, you must focus on short-term solutions instead of trying to correct long-term design issues. Immediate short-term solution is to try injecting some anti-foam chemical (unless a bunch has already been put in) and/or reducing the inlet gas rate. If these actions help, you can then track down the real cause and hopefully correct it. If anti-foam does not help, do not inject any more. Too much anti-foam can make things worse.

– It is important to note that the choice of an antifoam agent needs to be considered carefully. In general, silicone-based antifoam agent should be used in an amine-based solvent. The use of alcohol-based solvent is often preferred in hybrid or mixed solvent (mixture of physical solvent and amine solvent). Check with the amine vendor on their recommendation.

 

 

3.2         Troubleshooting Mol Sieve Dehydration Units (MSDU) [3, 4]

 

Figure 4: Example Process Flow Diagram for a Molecular Sieve Dehydration Unit [5]

 

Most operating problems that can occur in MSDU are:

– The adsorber produces specification product (i.e., water content at the outlet of the absorber) part of the time, but not for the entire cycle.

– The adsorber does not produce specification product at any time in the cycle.

– Pressure drop in the adsorber becomes so high that gas flow must be reduced due to lack of adequate pressure, or for fear of damage to internal bed support structure.

 

 

3.2.1     Low Water Adsorption Capacity on the Mol Sieve Bed

1. When the adsorption capacity is low, the plant produces specification outlet product during the first part of a cycle, and the product is off-spec during the latter part. When an adsorber indicates inadequate capacity, it is easy to decide that the adsorbent needs to be replaced. This could be the case, but a few checks should be made before investing in a new adsorbent bed.

A short-term solution could be to reduce the adsorption time but that reduction is limited to the time required to properly regenerate an exhausted adsorber bed.

2. Ensure the inlet gas conditions (P & T) have not changed, thereby increasing  water load entering the adsorber compared to design.  A relatively small increase in feed temperature can increase the water load significantly, typically, 1⁰C (1.8⁰F)  temperature increase of the gas would increase the gas water content by roughly 6%.

3. Other seemingly minor changes in operating conditions can have equally dramatic influence on the operation of an adsorber. A change in the wells, or formation from which the feed comes, can be important. Remember how various other components can compete for space on the adsorbent pore surface. Check to make sure the feed does not contain oxygenates (methanol), glycol or other chemicals that are injected routinely in operations.

4. Make certain that the meter used to measure the dew point or other contaminants is accurate. The test cells used in many modern instruments will occasionally pick up contamination which alters their level of response. It is good practice to keep a spare cell, or probe, and check the system periodically. Alternatively, one can consider installing two analyzers on the outlet.

5. Make certain that liquid hydrocarbons are not entering in the inlet gas to the adsorber. If this is the situation, the liquid hydrocarbons will coat the adsorbent and make the bed operate as if it were in a liquid system. In that case, the transfer of water from gas to adsorbent is very slow, thus the mass transfer zone is much longer than normal. This means there is less adsorbent to reach dynamic equilibrium and consequently the mass transfer zone is being extended. The result is a drastic decrease in capacity of the adsorber, and premature break-through.

6. Monitor the pressure drop across the bed. If there is a sudden increase in bed pressure drop, it can mean that some contaminant has entered the bed, or that the adsorbent beads or pellets have broken up. In either case the result may be poor flow distribution through the bed, and a resultant fast break-through. On the other hand, if the pressure drop suddenly falls quite a bit, it could mean that the adsorbent bed support has developed a leak and part of the adsorbent has been lost. This can be a significant problem in plants that do not have filters on the outlet of the adsorption plant. Typically, start-of-run ΔP would be 0.6 bar [8.7 psig] and end-of-run would be 1.2 bar [17.4 psig].

7. Make certain switching gas valves are not leaking. Feeling the outlet end of a closed valve during the heating phase is an easy way of detecting a leak.  If the piping is hot, the valve is leaking.

 

 

3.2.2     Mol Sieve Degradation [4]

Deactivation of a portion of the molecular sieve in natural gas dehydration units is inevitable. Typical degradation resulting in sieve deactivation include:

– Thermal degradation (is unavoidable)

– Caking (is avoidable)

– Coking (could be due to heavies such as heavy hydrocarbons, or gas treating solvents being adsorbed at the start-of run, thus difficult to avoid if present in the feed gas)

Deactivation can be more severe in the case of cake or coke formation. In time, the capacity at the top of the bed will be lower than at the bottom of the bed and this is often referred to as “capacity gradient”.

Let us look at more closely to these degradation phenomena.

 

 

3.2.2.1  Thermal Deactivation

With each thermal swing, the water adsorption capacity of the molecular sieve will decrease (thermal degradation). Thermal degradation will always occur.

Fines and dust formed during the regeneration cycle can also accumulate on the mercury removal bed downstream and result in an increased ΔP of that bed.  This can result in the possibility for retrograde condensation in the downstream bed.

How do you improve this? Optimise regeneration frequency (minimize number of regeneration cycles by applying variable cycling). At the start of the unit, there is adequate water adsorption capacity as the unit would be designed for end-of-life (i.e., lower) adsorption capacity. This means that longer adsorption cycle time could be used, thus minimising the number of regeneration cycles. This normally would extend the life of the mol sieve unit before a mol sieve change out is executed. A performance test run (PTR) is often used to determine the  degradation profile of the adsorbent bed.

 

 

3.2.2.2  Caking – Hydrothermal Deactivation

Caking is a term used to describe the loss of zeolitic and/or binder structure.

This can be caused by:

1. the presence of liquid water on the bed during regeneration (can be avoided by using a proper regeneration temperature profile).

2. liquid carry-over from an upstream unit. The cause of this is most often feed gas knock out drum upstream of the mol sieve bed being undersized or the mist extractor in the KO drum is not effective in catching fine liquid droplets that are being carried into the mol sieve bed.  Best practice is to use an adequately sized inlet coalescing filter separator rather than a conventional KO drum.

How do you improve this? Optimise the regeneration cycle i.e., slowly ramping up the temperature of the bed under regeneration to avoid formation of water on the top head / upper walls of the vessel during the heating cycle and check the KO drum sizing upstream of the MSDU.

 

 

3.2.2.3  Coke – Hydrocarbon Entrainment

Coking is a term used to describe blocking of micro and macro pores by (carbon) deposits. This is caused by the presence of ‘heavy’ components in the feed (e.g., amines, BTEX, C₂₀⁺). Once adsorbed during the adsorption cycle (this shall occur when the mol sieve is still fresh at start-of-run), these heavy components crack during the regeneration cycle causing coke on the mol sieve material. This often hinders the mass transfer phenomena and reduces the water uptake capacity on the sieve material.

How do you improve this? At the design stage, always think of installing a layer of guard material (silica-gel or alumina) on top of the mol sieve bed. Such a guard layer protects the mol sieve bed by “catching” impurity droplets upstream of the mol sieve bed.  This must be done during the design stage so that there is adequate vessel capacity to hold the required mass of the sieve and the guard material layer.

 

 

4.0         Troubleshooting Mercury Removal Unit (MRU)

Most common adsorbents used for the removal of metallic mercury are:

• Sulphur impregnated carbon:

• Metal Sulphided Adsorbents consisting of CuS and ZnS or PbS:

If early breakthrough of mercury is detected, the most likely causes to investigate are:

– Check if the mercury analyzer has calibration issue.

– Check if the mercury content in the feed gas to the LNG / Gas Plant has changed due to the change in the upstream operating wells.

– Check the phase envelop of the feed gas entering the MRU. Is there possible retrograde condensation or any liquid fall out on the adsorbent bed? If yes, it is possible that the sulphur that is impregnated on the activated carbon has leached out causing the adsorbent to be inactive in mercury capture. If a metal sulphided adsorbent is used, then the liquid fall out is causing slower mass transfer rates and both saturation and mass transfer zones are much longer than the design case, causing an early breakthrough of mercury at the outlet of the MRU.

There is new adsorbent on the market that has been developed by Queen’s University’s Ionic Liquids Laboratory (QUILL) and Petroliam Nasional Berhad (PETRONAS). The process, marketed as HycaPureHg™, captures all mercury species present in natural gas and the researchers claim that it has up to 3 times higher capacity than competing state-of-the-art commercial alternatives.

The jury is still out as we await publication of the results on the trials carried out on Malaysian LNG plants.

 

 

5.0         Concluding Remarks

In this TOTM, we discussed some common operational issues of the three critical process units in LNG and Cryogenic Gas Processing Facilities.

Troubleshooting these units requires a good understanding of the feed gas components and the possible contaminants that would affect optimal operation of these units.

If the reader is interested further on these subjects, we suggest attending our G4 (Gas Conditioning and Processing); G4 with LNG Emphasis, G6 Gas Treating and Sulphur Recovery, or our new 3-day short courses on Acid Gas Removal Fundamentals, and Molecular Sieve Dehydration Fundamentals courses provided by PetroSkills/JM Campbell.

 

 

References:

1. The Basics of Gas Treatment for Liquefaction Plant by J. Castel, Paper presented at the GPAE held in Hamburg, Germany, 23^rd April 2015

2. Gas Conditioning and Processing, Volume 4: Gas Treating and Sulphur Processes by R.N. Maddox and D. John Morgan, John M. Campbell Publications

3. Introduction to Oil and Gas Productions Systems, Version: BP-IEP_01.03.2004

4. Finding the Fountain of Youth for a Mol Sieve Dehydration Unit by A.F. Carlsson
J.B. Rajani, and A.J. Kodde, Paper presented at the 83^rd GPA convention held in
New Orleans, USA, March 2004

5. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^th Edition, 3^rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

Abstract

The construction costs of LNG storage tanks are very high, mainly due to the availability and cost of nickel. Most LNG storage tanks built in the modern era use 9% wt nickel steel as the materials of construction because of the materials well documented toughness at cryogenic temperatures. This article looks at advances in cryogenic metallurgy in order to reduce construction costs while maintaining a safe operational environment.

 

Introduction

The May 2021 Tip of The Month (TOTM) discussed the rationale behind the choice of 9% wt Ni steel for inner tank construction of aboveground LNG storage tanks. This TOTM explores the development of 7% wt Ni steel for the same application with the potential for significant cost savings while maintaining a high level of toughness and strength at LNG temperatures.

The advantages of 9% steel at cryogenic temperatures are well documented and the steel is specified in API 620 Appendix Q as ASTM A553 type 1. The ASME Boiler & Pressure Vessel Code Part II has a chemical requirement of SA-553 Type 1 (equivalent to ANSI A553 Type 1) that requires a nickel content of 8.50% to 9.50% with a tensile strength of 690 MPa to 825 MPa (100,050 to 119,625 psi). In addition, the critical parameter is that the material must have a toughness such that a longitudinal Charpy V-notch test energy shall be not less than 34 Joules (0.0323 BTU) at the specified temperature while a transverse V-notch energy shall be not less than 27Joules (0.0257 BTU). The mechanical properties of SA-553 make it a good option for operation at very low temperatures at low stresses.

There is an opportunity for significant cost savings, though, if the nickel content could be safely reduced. With the price of nickel heading towards US$20,000 per tonne, even a small reduction in nickel content could bring a significant cost reduction in the price of a LNG storage tank. Table 1 below shows the projected cost savings of changing to 7% wt nickel instead of 9% wt nickel for three different sized tanks with a nickel cost of US$17,500/tonne.

 

Table 1 Nickel cost savings in LNG storage tanks of various sizes (thanks to Dr Jay Rajani)

 

LNG storage tanks have used 9% nickel steel as the main material for the last fifty years; it has been the industry standard for LNG tank material since it was developed by the International Nickel Company in the USA in 1946.  An evaluation of this steel (Operation Cryogenics) showed that stress relief was unnecessary for 9% nickel steel making it much cheaper to manufacture than contemporary high nickel steels.  The difficulty in using even lower quantities of nickel is in attaining the same degree of metal toughness. Maintaining the required toughness and strength for cryogenic applications has been the primary barrier to further nickel content reduction in the steel for many years.

The nickel content of SA-553 steel gives excellent cryogenic fracture toughness due to retained austenite and the fine microstructure from the nickel content combined with specific heat treatment processes.  In the past, reducing the nickel content and using the same heat treatment resulted in a reduction in cryogenic toughness that is unacceptable in LNG storage applications.

Around 1960 the Nippon Steel and Sumitomo Metal Corporation (NSSMC) initiated a research and development program to investigate nickel content while maintaining cryogenic toughness.  The objective of the development program was to develop a steel that resisted both crack initiation and propagation. The steel production method that was used was TMCP (thermo-mechanical controlled processing).

The processes involved in TMCP are:

• Thermo-Mechanical Rolling (TMR)
  • a technique designed to improve the mechanical properties of materials by controlling the hot-deformation process in a rolling mill
• Accelerated Cooling (AC)
  1. Makes it possible to significantly increase strength properties without lowering impact toughness or cold resistance.
  2. Makes it possible to replace the ferritic-pearlitic structure usually formed in steel after conventional controlled rolling with a fine-grained ferrite-bainite structure having a diminished level of striation.
  3. Makes it possible to attain a prescribed level of strength with lower quantities of carbon and alloying elements. That in turn improves the weldability of the steel.
  4. Reduces the load on the mill because a higher finishing temperature is used than in traditional CR (controlled rolling), allowing an increase in rolling speeds through a reduction in the number of pauses made to cool the slabs on the mill.
• Direct Quenching & Tempering (DQ&T)
  • This approach of direct quenching uses plate recalescence (the surface reheating following accelerated cooling due to the heat flow from its core that is still hot) to promote a direct tempering to the product.

 

Both TMR and AC had previously been used in shipbuilding steel and line-pipe manufacture with good results.  DQ&T with a low heating temperature was shown to reduce crack propagation and improved mechanical properties.  A further process (Lamellarizing) is used to form reverted austenite, improving impact toughness.

Figures 1 and 2 show the difference in heat treatment from conventional heat treatment to the refined heat treatment employed in the thermo-mechanical control process.

 

Figure 1 Conventional Heat Treat for 9% wt Ni Steel [1]

 

 

Figure 2 TMCP Heat Treat process used in 7% wt Ni Steel [1]

 

Table 2 below shows the resultant mechanical properties of A553 Type 1 (9% wt Ni steel) and A841 Grade G, Class 9 (TMCP 7% wt Ni steel) that has been manufactured using the TMCP process.  Notice that the tensile strength (TS), the ductility (elongation – El) and, crucially, the Charpy V-notch (IV) energy are all identical.

 

Table 2 Properties of A553 Type 1 and A841 Grade G, Class 9

Specification	ASTM
	A553 Type I	A841 Grade G
		Class 9
Plate thickness (mm)	50 max.	50 max.
Process	QT	TMCP
C (%)	0.13 max.	0.13 max.
Si (%)	0.15-0.40	0.04-0.15
Mn (%)/td>	0.90 max.	0.60-1.20
P (%)	0.035 max.	0.015 max.
S (%)	0.035 max.	0.015 max.
Ni (%)	8.50-9.50	6.00-7.50
0.2%PS (MPa)	585 min.	585 min.
TS (MPa)	690-825	690-825
El (%); Thick(mm)	20 min.	20 min.
IV (J) at -196°C	34 min.	34 min.
LE*1 (mm) at -196°C	0.38 min.	0.38 min. (t ≤ 32)
		0.48 min. (t=50)*2
Note: *1: LE: Lateral Expansion
*2: LE value between the plate thickness 32 and 50 shall be determined　by linear interpolation.

Source: https://www.kobelco-welding.jp/education-center/technical-highlight/vol15.html

 

This variant of nickel steel has been approved as a Japanese standard under JIS G 3127, SL7N 590 and has been used successfully in LNG tank construction in Japan and the Middle East.  The second addendum to API 620 twelfth edition Annex Q (2018) now allows the use of A553 Type III (reference ASME Code Case 2842 – specifically for pressure vessels) or A841 Grade G, Class 9 as a standard material for product temperature applications down to -325ºF (-198ºC).  The use of A841 Grade G, Class 9 requires that SA-841 Supplementary Requirement S64 has been satisfied.  S64 is defined in ASME Boiler & Pressure Vessel Code Part II as follows:

S64.1 Except for the TMR-I-T process, the plates shall be cooled directly after rolling without being allowed to cool below 1025°F [550°C]. Quenching hardening shall be initiated from a temperature within the range from 1025 to 1490°F [550 to 810°C].

S64.2 Subsequent to quenching, the plates shall be tempered within the range from 1030 to 1155°F [555 to 625°C], holding at that temperature for a minimum of 30 min/in. [1.2 min/mm] of thickness but for not less than 15 min, and then cooling at a rate of not less than 300°F/h [165°C/h], either in air or by quenching in water, to ambient temperature.

S64.3 Prior to the tempering treatment, the plates may be subjected to an intermediate heat treatment (Note S64.1) consisting of heating to a temperature in the range from 1185 to 1310°F [640 to 710°C], holding at that temperature for a minimum of 1 hr/in. [2.4 min/mm] of thickness, but in no case less than 15 min, and then water-quenching to below 300°F [150°C] in the case of plate thicknesses of more than 5⁄8 in. [16 mm]; or cooling in air or water-quenching in the case of plate thickness of 5⁄8 in. [16 mm] and under.

NOTE S64.1—The intermediate heat treatment is for the purpose of enhancing elongation and notch-toughness and for reducing susceptibility to strain-aging embrittlement and temper embrittlement. It may be performed at the option of the material manufacturer or may be specified by the purchaser.

S64.4 Heat treatment temperatures and times shall be reported in accordance with Section 19 of Specification A20/A20M.

The potential for cost savings runs into hundreds of thousands of US dollars for each LNG tank meaning this material should be considered as a valid alternative to 9% nickel steel for future LNG inner tank construction.

 

 

References

1. Hiroshi Nagami et al., Development and realization of large scale LNG storage tank applying 7% Nickel steel plate, Kuala Lumpur World Gas Conference 2012
2. Hitoshi Furuya et al., DEVELOPMENT OF LOW-NICKEL STEEL FOR LNG STORAGE TANKS, Nippon Steel & Sumitomo Metal Corporation
3. Takayuki Kagaya et al., New Steel Plate for LNG Storage Tank, Nippon Steel & Sumitomo Metal Corporation technical report No. 110 September 2015.
4. Antonio Augusto Gorni et al., Accelerated Cooling of Steel Plates: The Time Has Come, Journal of ASTM International, Vol. 5, No. 8, Paper ID JAI101777,  Available online at www.astm.org
5. API 620 Design and Construction of Large, Welded, Low-Pressure Storage Tanks, Twelfth Edition, October 2013, Addendum 2, April 2018
6. ASME Boiler and Pressure Vessel Code, Section II, Part A
7. ASME Code Cases July1, 2019.  CC2842

The base material for the tank containing the liquid gas (such as LNG) at below -165°C (-265°F) must remain ductile and crack resistant with the highest level of safety. The material must also permit welding without any risk of defects, for example, induced brittle fracture. Stainless steels, aluminum and 9% nickel steels can be used as they do not have a ductile/brittle transition temperature. However, in practice aluminum and stainless steel have become uneconomic for large land-based tanks but aluminum alloys are used for the large spherical tanks in gas tankers because of the lower weight. 9% nickel steel provides an attractive combination of properties at a moderate price. A high corrosion resistance is not required for LNG tanks.

Steels alloyed with nickel are used in many cryogenic applications since nickel improves the quench ability and improves the notch toughness at low temperatures. Steels with 3.5% nickel, 5% nickel and 9% nickel are used at temperatures below -50°C (-58°F). At temperatures below -104°C down to -196°C (-155°F down to -320.8°F) mainly the 9% nickel steels are used. The 9% nickel steel was developed in the early 40s following the “The Disaster of the Cleveland East Ohio Gas Explosion” in 1944.

A new LNG tank was manufactured with low nickel content (approximately 3.5%) because stainless steel alloys were limited in availability due to World War II. On October 20, 1944, the tanks had been filled to their capacity in readiness for the coming winter months. At some time during the afternoon, regular mechanical stress on the fourth tank’s walls created a leak, which lead to failure of the outer carbon steel shell, causing the LNG tank to completely fail and thus discharging all its contents into the nearby streets and sewers of Cleveland. The LNG evaporated quickly and formed a flammable gas cloud. The liquid LNG flowed into the gutters and storm sewers. A gas cloud eventually found an ignition source creating a massive explosion.

Since this incident, 9% Ni become the steel for LNG tanks. For lower temperatures (liquid hydrogen minus 252.8°C or minus 423°F) stainless steels are used for vessels.

 

Lower nickel content steel development for LNG containment tanks

Steel with nickel levels from 6 to 9% is able to remain ductile and crack resistant when containing materials of extremely low temperatures, including LNG. This temperature resistance is the result of the fine-grained structure of nickel-ferrite. Heat treatment makes this temperature resistance even more effective and effectively making these alloys very strong. Steel nickel alloys can be created with reduced wall thickness, creating cost-efficient solutions for constructing LNG storage and transportation tanks. Other parameters that play a role in the choice of material is the grain structure of the steel, brittleness, weldability of the steel and plastic deformation.

It is also essential for LNG tanks to be constructed from materials that have excellent weldability. Nickel alloy levels from 6% to 9% are well known for their weldable properties and traits. Another important issue related to welding is temperature resistance. Welded joints need to maintain the highest levels of toughness and resilience. Otherwise, brittle fracture could result, allowing for the possible vaporization of the LNG. Fortunately, a great deal of research has been done to establish the safest materials for welding LNG tanks. Nickel alloys have withstood temperatures of -196°C (-320.8°F) without losing its toughness [1, 2].

However, for stainless steel to have the best resistance to cracking, there are certain elements that need to be present. A minimum level of delta ferrite (this is a high temperature form of iron, formed on cooling low carbon concentrations in iron-alloy from the liquid state before transforming to austenite) is one. Stainless steel with higher levels of ferrite reduces the impact of low temperatures. When applied with sub-arc flux and covered electrode coatings, the stainless steel maintains a lower amount of micro-slag and higher ductility as a result. Stainless steel with these qualities makes an ideal partner for nickel in the construction of LNG tanks and containers.

Kobelco of Japan [3] has given a good graph, shown in Figure 1, that shows the applicability of the various steels for cryogenic application.

 

Figure 1. The boiling points of various liquefied gases and applicable metals for storage tanks [3]

 

A wide range of regulations and standards define the design, construction, inspection, and maintenance of LNG tanks made of 9% Ni steel. Some of the relevant ASME, API, BS EN, and JIS standards are provided below.

1. ASME Sec. VIII, Div. 1: Design and fabrication of pressure vessels; Div. 2: Alternative rules.
2. API Standard 620: Design and construction of large, welded, low-pressure storage tanks; Appendix Q: Low pressure storage tanks for liquefied hydrocarbon gases at temperature not lower than −270°F (−168°C).
3. BS EN 14620-1(2006): Design and manufacture of site built, vertical, cylindrical, flat-bottomed steel tanks for the storage of refrigerated, liquefied gases with operating temperatures between 0°C and −165°C (−265°F), Part 1: General.
4. JIS B8265(2010): Construction of pressure vessel ― General principles; JIS B8267(2008): Construction of pressure vessel.

One way of looking [4] at the viability of using lower nickel content steels is to use impact testing by the use of Charpy Impact vs Test Temperature Test. The Charpy Impact Test, also known as the Charpy V-notch test, is a standardized high strain-rate test which determines the amount of energy absorbed by a material during fracture. Generally, low alloy nickel steels show rapid falloff of the Charpy impact resistance at low temperatures. Since the Ohio incident, rigorous laboratory tests showed that 3.5% Ni steel was not adequate for LNG temperatures.

Another aspect that needs to be looked at is the steep rise in toughness even at very low temperature, as the nickel content of the alloy is increased to between 8 to 9%. This increased toughness shows that Charpy values well above 15 ft-lbs (20 Nm) minimum required by the ASME Boiler code could be met by these alloys. This search for a tough alloy that could be used at low temperatures and produced at reasonable cost seems to have paid off. This material, 9%Ni steel, eventually came to be the material of choice for LNG storage tanks.
Summary of Discussion

For LNG storage tanks, 9%Ni has proven to be a safe material of choice. Remember that the LNG storage tanks have an outer containment concrete shell and inner 9%Ni tank at atmospheric pressure. For LNG piping (no containment provision is provided), however, 36%NiFe is the preferred choice due to the diverse mechanical and physical properties such as low coefficient of expansion. The stress is lowest in 36%NiFe because of its extremely low coefficient of expansion, whilst the stress is high in 9%Ni, its high strength prevents this alloy from yielding because of the thermal strain of contraction.

For LNG storage tank, the move to 7%Ni is purely due to cost reduction. However, the application of 7%Ni needs to have various approvals from certifying authorities (AMSE, Lloyds Register, ABS, DNV-GL, etc.).

DNV-GL of Norway has not approved 7% Ni steel for LNG application. However, the IGC code (Table 6.3) provide an opening to use 5% Ni steel for temperatures down to −165°C or −265°F (LNG). It will accordingly be possible to also accept 7% Ni Steel on the same basis, i.e., provided impact testing at -196°C (−320.8°F) show acceptable results (other tests may also be required based on the material properties and the chemical composition).

If you are interested in LNG Shipping, LNG Storage, LNG Receiving Facilities and LNG Safety, we suggest attending our G4 (Gas Conditioning and Processing) or G4 with LNG Emphasis courses provided by PetroSkills/JM Campbell.

By: Dr. Jay Rajani, Senior PetroSkills Instructor
References
1.    Development of Low-Nickel Steel for LNG Storage Tanks by Hitoshi Furuya et al of Nippon Steel and Sumitomo Metal Corporation.

Click to access Materials-2-Hitoshi_Furuya-LNG17-Poster.pdf

2.    New Steel Plate for LNG Storage Tank by Takayuki Kagaya et al
NIPPON STEEL & SUMITOMO METAL TECHNICAL REPORT No. 110 SEPTEMBER 2015
3.    9% Ni Steel is commonly used for above ground LNG tanks.
https://www.kobelco-welding.jp/education-center/technical-highlight/vol02.html
4.    Liquified Natural Gas in the United States: A History by John Hrastar

The solubility of hydrocarbons and non-hydrocarbons like carbon dioxide and hydrogen sulfide in water is of interest for oil and gas production and processing facilities, dealing with water treatment, and disposal facilities from an environmental aspect. Understanding the water solubility of gases in produced water is critical for piping and equipment sizing to ensure there is adequate vapor handling capacity, for example, in the production separators, gas compressors, vapor recovery units, or possibly thermal oxidizers. This is important to ensure the facilities have adequate capacity to minimize hydrocarbon emissions. In addition, the determination of the solubility of these components in aqueous phase is critical in the study of the gas hydrates kinetics [1].

In Parts 1 and 2 of this series, January, and March 2021 tips of the month (TOTM) [2, 3], we focused on the solubility of light hydrocarbon gases such as CH₄, C₂H₆, C₃H₈, iC₄H₁₀, and nC₄H_10­. We reviewed and presented an example of available experimental solubility data, a thermodynamic model. For that case study, we presented diagrams to show the effect of temperature and pressure on the solubility of these light hydrocarbon gases in water.

We also presented examples for the following applications:

1. Converting the gas solubility from std m³ of gas per m³ of water (scf of gas per bbl of water) to mol fraction.

2. A shortcut method for estimating gas solubility in water.

3. Using the solubility charts with compositions of Gas A to estimate the solubility of another gas such as Gas B with different compositions.

4. Converting the gas solubility from mol fraction to std m³ of gas per m³ of water (scf of gas per bbl of water).

 

Table 1. Compositions of two gas mixtures

 

 

Case Study: How do I size the water tank vent line to VRU compressor?

a. Process simulation

To investigate the estimation of the vent line size of the water tank, we will consider the production process flow diagram of Figure 1 (FPS). The SI version of Figure 1 (SI) is in Appendix A. the simulation results are obtained from ProMax [4] using the Soave-Redlich-Kwong equation of state.

 

(click to enlarge)

 

Figure 1(FPS). Case study process flow diagram simulation results by ProMax [4]

 

 

The oil compositions and the heavy end components properties are presented in Table 2.

Table 2. Case study oil compositions and the heavy ends properties

 

As shown in Figure 1, 5000 stock tank bbl/day (33.122 STm³/h) of oil with the compositions shown in Table 2 is mixed with 5000 bbl/day (33.122 m³/h) of water at 80 °F and 120 psia (26.8 °C, 827.4 kPa) enters the three-phase separator (VSL-100). The separated water containing the dissolved and entrained hydrocarbons (stream 6) enters the gun barrel from which the separated water (Stream 9) goes to the water tank. Finally, the water (Stream 11) at a rate of 4,998.5 bbl/day (33.113 m³/h) goes to the salt-water disposal (SWD) facilities.  The tank vapor stream at the rate of 400.49 SCF/day (11.3406 Sm³/d) goes the VRU system. The solution gas, R_S (Stream 9 to water tank) is 400.49 scf/4,998.5 bbl or R_S = 0.0801 scf/bbl (R_S =11.3406/33.113×24 = 0.01427 scm/m³).

 

a. Vapor recovery unit (VRU)

Figure 2 presents the schematic of water tank and its VRU system. The VRU Compressor and Scrubber have already been evaluated and are suitable for the forecasted production.  The vapor recovery line connecting the tank to the VRU compressor has not yet been evaluated. What size Vapor Recovery Line is adequate to handle the vapors generated by the inflow of water?

 

Figure 2. The schematic of water tank and its VRU system

 

Production Conditions

Water Rate: 6,000 STBWPD (954 STm³/d)

The solution gas in feed to water tank is R_S = 0.0801 scf/bbl of water (R_S = 0.01427 scm/m³)

Upstream Gunbarrel Operating Pressure: 5 psig (35 kPag)

Upstream Separator Operating Temperature: 80°F (27°C)

►It is common to allow the water level in the Water Storage Tank to build up over time, thus the pump out rate from the Water Tank to the Salt Water Disposal (SWD) facilities is often intermittent.

►To determine the VRU line size, assume that the tank is in filling mode, thus the pump out rate from the tank is zero, but production from the upstream separator continues.  Total gas flow is then equal to the volume of water entering the tank (pushing the same volume of the gas out of tank), plus the gas that will evolve from the produced water (0.0801 scf/bbl or 0.01427 scm/m³).

►For the VRU line size, we also need to consider the case when water rate to SWD is the same as production from the upstream separator. Total gas flow is then equal to the volume of the gas that will evolve from the produced water (0.0801 scf/bbl or 0.01427 scm/m³)).

►During tank pump-out operations, the blanket gas system will be designed to accommodate the water removal rate from the tank, which is a different calculation than the VRU line sizing

Spitzglass formula is:

 

 

 

 

►Rearranging to solve for head loss (Δh_w):

 

FPS SOLUTION

Assume that the water flowrate is equivalent to Stock Tank flow (the temperature (80°F) is slightly higher than ‘stock tank’ (60°F). (i.e. assume Q_water = 6000 bbl/d)

Q_G = Qwater + Qwater*Rs, where Rs is released gas.

Therefore, gas evolving from the tank due to full water rate but zero output while the tank is filling (water tank pumps are not running):

Q_G = 6,000(bbl/d) x5.615(scf/bbl) +6,000(bbl/d) x 0.0801(scf/bbl) = 34,171 scfd = 0.034171 MMscfd

Let us try, ID= 2.469 inches (62.7 mm), L = 600 ft (182.9 m), and API 2000 [5] recommends S = 1.5; therefore, this value was used even though the estimated gas relative density was S=0.84.

 

Substituting in:

 

 

Assuming the pressure at the VRU compressor suction is 1 oz/in², the pressure at the tank would have to be 3.3 oz/in².  This is less than the design pressure of the tank, and it is also less than the PVSV opening pressure (set at 3.5 oz/in²).

Next check if the system will handle the gas assuming that the water flow rate out of the tank is equal to the water flow rate coming into the tank.  If (Q_W)_IN = (Q_W)_OUT, then the gas rate is a result of gas coming out of solution so:

Q_G =R_S * Q_W =0.0801 scf/bbl x 6,000 bbl/d = 481 scfd = 0.000481 MMscfd

 

Substituting back into Spitzglass yields:

 

This pressure is too small to be of concern. If the pressure at the VRU compressor is 1 oz/in², then pressure at the tank would be slightly more than 1 oz/in².  This is less than the set point of the flare valve (2 oz/in²), and less than PVSV opening pressure (set at 3.5 oz/in²). The VRU compressor is going to run at very low rates frequently and the compressor vendor would need to verify that the selected compressor can operate at all anticipated operating conditions of the system

 

 

SI SOLUTION

Assume that the water flowrate is equivalent to Stock Tank flow (the temperature (27°C) is slightly higher than ‘stock tank’ (15°C). (i.e. assume Q_water = 954 m³/d).

Q_G = Qwater + Qwater*Rs, where Rs is released gas.

Therefore, gas evolving from the tank due to full water rate but zero output while the tank is filling (water tank pumps are not running):

Q_G = 954 m³/d + 954 m³/d x 0.01427 scm/m³= 968scm/d = 0.000968×10⁶ scm/d

Let us try, ID= 2.469 inches (62.7 mm), L = 600 ft (182.9 m), and API 2000 [5] recommends S = 1.5; therefore, this value was used even though the estimated gas relative density was S=0.84.

Substituting in:

 

 

Assuming the pressure at the VRU compressor suction is 431 Pa (1 oz/in²), the pressure at the tank would have to be 1.422 kPag (3.3 oz/in²).  This is less than the design pressure of the tank, and it is also less than the PVSV opening pressure (set at 1.509 kPag = 3.5 oz/in²).

Next, check if the system will handle the gas assuming that the water flow rate out of the tank is equal to the water flow rate coming into the tank.   If (Q_W)_IN = (Q_W)_OUT, then the gas rate is a result of gas coming out of solution so:

Q_G =R_S * Q_W =954 m³/d x 0.01427 scm/m³= 13.6136 scm/d = 0.000014×10⁶ scm/d

 

Substituting back into Spitzglass yields:

 

 

This pressure is too small to be of concern. If the pressure at the VRU compressor is 431 Pa (1 oz/in²), then pressure at the tank would be slightly more than 1 oz/in².  This is less than the set point of the flare valve (862 Pag = 2 oz/in²), and less than PVSV opening pressure (set at set at 1.508 kPag = 3.5 oz/in²). The VRU compressor is going to run at very low rates frequently and the compressor vendor would need to verify that the selected compressor can operate at all anticipated operating conditions of the system.

Table 3 shows the comparison of pressure drops estimated by three correlations and indicates that the Spitzglass value is the most conservative.

 

Table 3. Vent line pressure drop by three correlations

 

Summary

This TOTM presented the computer simulation results for an oil production facility (Figure 1) with a watercut of 50% by volume. From the simulation result, the produced water rate and the vent gas rate from the water tank were used to estimate the diameter of the vent pipe to the VRU compressor (Figure 2). The diameter was estimated by the Spitzglass correlation under pressure setting constrained shown in Figure 2. In addition, the vent line pressure drop was estimated by Weymouth and Oliphant correlations (Table 3). Table 3 indicates that the Spitzglass value is the most conservative. This is important because often, these systems operate with very tight tolerances.  Notice that when the saltwater disposal system is not removing water from the tank (our design case here), the pressure in the tank is 3.3 oz/in² [1.422 kPag] and the opening pressure for the PVSV is 3.5 oz/in² [1.422 kPag].  This provides very little margin for error, justifying the use of the most conservative algorithm (e.g. Spitzglass).  It also highlights the importance of designing this vapor line with ample drainage, and “No Pockets” or traps where liquids can accumulate.

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, and Applications in Gas Processing) courses.

 

By: Mahmood Moshfeghian, Ph.D., & Kindra Snow-McGregor, P.E.

 

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References:

1. Sloan, E.D., Koh, C.A., Clathrate Hydrates of Natural Gases, 3^rd ed., CRC Press, 2008.

2. Moshfeghian, M., http://www.jmcampbell.com/tip-of-the-month/2021/01/solubility-of-gases-in-water-part-1/, Light Hydrocarbon gases, Jan 2021

3. Moshfeghian, M., https://www.petroskills.com/blog/entry/00_totm/mar21-fac-Solubility-of-Gases-in-WaterP2#.YFoq4VVKipo , Mar 2021

4. ProMax 5.0, Build 5.0.20034.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2020.

5. API Standard 2000, Venting Atmospheric and Low-Pressure Storage Tanks, 7^th Edition, American Petroleum Institute. 2014

6. GPSA Engineering Data Book, Section 17, Volume 2, 13^th Edition, Gas Processors and Suppliers Association, Tulsa, Oklahoma, 2012.

7. Campbell, J.M., “Gas Conditioning and Processing, Volume 1: The Fundamentals,” 9^th Edition, 3^rd Printing, Editors Hubbard, R. and Snow–McGregor, K., Campbell Petroleum Series, Norman, Oklahoma, PetroSkills 2018

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Appendix A

Vent gas rate is Q_G = 34171 scfd (From FPS Solution section)

Oliphant Correlation [6]

 

 

Weymouth Correlation [7]

 

 

 

Figure 1(SI). Case study process flow diagram simulation results by ProMax [4]

Why should we care about solubility of gases in water?

The solubility of hydrocarbons and non-hydrocarbons like carbon dioxide and hydrogen sulfide in water is of interest for oil and gas production and processing facilities. dealing with water treatment, and disposal facilities from an environmental aspect. Understanding the water solubility of gases in produced water is critical for piping and equipment sizing to ensure there is adequate vapor handling capacity, for example, in the production separators, gas compressors, vapor recovery units or possibly thermal oxidizers. This is important to ensure the facilities have adequate capacity to minimize hydrocarbon emissions. In addition, determination of the solubility of these components in aqueous phase is critical in the study of the gas hydrates kinetics [1].

In Part 1 of this series, January 2021 tip of the month (TOTM) [2], we focused on the solubility of light hydrocarbon gases such as CH₄, C₂H₆, C₃H₈, iC₄H₁₀, and nC₄H_10­. We reviewed and presented example available experimental solubility data, example thermodynamic model. For that case study, we presented diagrams to show the effect of temperature and pressure on the solubility of these light hydrocarbon gases in water.

 

Case Study

To investigate the estimation of solubility of light hydrocarbon gases in water, we will consider two lean gas mixtures with the composition shown in Table 1.

 

Table 1. Compositions of two gas mixtures

 

 

In Part 1, for Gas A, we obtained the solubility of each component in water from ProMax [3] using the Peng-Robinson equation of state. Four isotherms of 5°C, 10°C, 20°C, and 30°C (41°F, 50°F, 68°F, and 86°F) were selected. For each isotherm, pressure was varied from 200 kPa to 8000 kPa (29 psia to 1160 psia). The calculated gas solubility results, std m³ of gas per m³ of water (SCF of gas per bbl of water) are presented in Figures 2 and 3 (See Appendix A).

 

In this tip, we will review and answer the following questions:

1. How do I convert the gas solubility from std m³ of gas per m³ of water (scf of gas per bbl of water) to mol fraction?

2. Is there a short cut method for estimating gas solubility in water and how can I use it?

3. How can I use the solubility charts of Figures 2 and 3 for compositions of Gas A to estimate solubility of another gas such as Gas B with different compositions?

4. How do I convert the gas solubility from mol fraction to std m³ of gas per m³ of water (scf of gas per bbl of water)?

 

 

Q1: Example 1 – Conversion of solubility units from volumetric rates to mol %

From Figure 3 of Part 1, for temperature of 30 °C (86 °F) and pressure of 6 MPa (870 psia) the solubility of gases in volumetric units are:

 

Convert gas solubility units from volumetric rates to mol %. Density of water is 1000 kg/m³ (62.4 lbm/ ft³) and MW = 18

 

SI Solution:

Assume a basis of one m³ of liquid water

 

 

Repeat the above calculation for each gas component and populate the second column from right of Table 2 SI. We also add 55.5556 kmol of water. The sum is 55.6161 kmol.

Repeat the above calculation for each gas component and populate the last column of Table 2 SI.

Table 2 SI. Conversion of volumetric solubility to mol%

 

 

FPS Solution:

Assume a basis of one bbl of liquid water

 

 

Repeat the above calculation for each gas component and populate the second column from right of Table 2 FPS. We also add 19.4653 lbmol of water. The sum is 19.4865 lbmols.

Repeat the above calculation for each gas component and populate the last column of Table 2 FPS.

 

 

Table 2 FPS. Conversion of volumetric solubility units to mol%

 

 

Q2: Solubility of light Hydrocarbon Gases – Henry’s Law

Henry’s law is one of the methods for gas solubility when the concentration of components in the liquid phase is low and the assumption of ideal gas is valid at low pressure. Under these conditions Henry’s law can be written by Equation 1,

Py_i=H_ix_i

(1)       Where P, H_i, y_i and x_i are the pressure, Henry’s constant, and mol fraction in the vapor and liquid phases of component i, respectively. For a mixture at high pressure, the non-ideality of the gas phase is described by the vapor phase fugacity coefficient, ϕ_i^V  in Equation 2.

 

Py_iϕ_i^V=H_ix_i

(2) ϕ_i^V should be calculated at prevailing conditions by an equation of state.

Henry’s constant of a component is experimentally determined. Also, it is assumed to be independent of its concentration but, this constant is a function of temperature. Figure 1 shows Henry’s constant for gaseous components of reservoir fluids in water [4].

 

 

 

Figure 1. Henry’s constant for different gases [4]

 

 

In addition, Henry’s constant dependency on pressure can be calculated using the following relation:

(3)

 

Where H_i⁰  and v_i∞  are Henry’s constant at P⁰ and the partial molar volume of component i in the solvent at infinite dilution. v_i∞  is assumed to be constant over the prevailing composition and pressure ranges. Equation 3 is known as Krichevsky-Kazarnovsky equation. Average of partial molar volume for methane, ethane, propane and nitrogen are 0.040, 0.055, 0.080 and 0.035 m³/kmol (0.641, 0.881, 1.281, 0.561 ft³/lbmol), respectively [4].

 

 

Q2: Example 2                     

For the gas mixture A of Table 1, estimate the methane, ethane, and propane gas solubility in water using Henry’s law at the following conditions: T= 30 °C (86 °F) and P = 8 MPa (1160 psia)

Compare the results with the ProMax results reported in Part 1.

 

Solution:

 

 

 

The solubility of each gas component can be calculated using Equation (2).

  (2)

 

Because of the low volatility of water in comparison to methane: y_C1=0.85. Also, from Table 1A in Appendix A, the methane fugacity coefficient is ϕ_C1^V=0.896. The gas solubility is as follows:

 

Based on example 1, the ProMax value of solubility is 0.1223%.

 

From Table 1A in Appendix A, the ethane fugacity coefficient is ϕ_C2^V=0.606.

Based on example 1, the ProMax value of solubility is 0.0104%.

 

From Table 1A in Appendix A, the propane fugacity coefficient is ϕ_C3^V=0.442.

 

 

Based on example 1, the ProMax value of solubility is 0.0025 %.

 

In this example, the accuracy of results has been improved by calculating the ϕ_i^V using an equation of state.

Table 2. Example 2 Summary of results for GAS A and comparison with ProMax results

 

 

 

Q3: Example 3

For the gas mixture B of Table 1, estimate the methane, ethane, and propane gas solubility in water at T= 30 °C (86 °F) and P= 7 MPa (1015 psia)

a. Using results of ProMax for gas mixture A at T= 30 °C (86 °F) and P= 8 MPa (1160 psia) listed in Table 2 and Equations 3 and 4

          (4)

Subscript 1 represents Gas A and its conditions and properties and subscript 2 represents Gas B and its conditions and properties.

b. Using Equations 2 and 3

c. Using of ProMax results for gas mixture B at T= 30 °C (86 °F) and P= 7 MPa (1015 psia) in Table 2A SI of Appendix A.

 

 

Solution

Method a:

From example 2, for gas mixture A at T= 30 °C (86 °F) and P₁= 8 MPa (1160 psia)

 

 

 

Using Equation 4 and fugacity coefficients from Table 1A of Appendix A:

 

 

 

Method b.

For gas mixture B at T= 30 °C (86 °F) and P₂= 7 MPa (1015 psia) from Method a.

 

 

 

Using Equation 2 and fugacity coefficients from Table 1A of Appendix A, the gas solubility for each component is as follows:

 

 

Method c.

For gas mixture B at T= 30 °C (86 °F) and P₂= 7 MPa (1015 psia) from Table 2A SI of Appendix A.

x_2C1=0.1155  mol%,  x_2C2=0.0089 mol%, x_2C3=0.0020 mol%

 

Table 3. Example 3 Summary of results for GAS B and comparison with ProMax results

 

 

Q4: Example 4 – Conversion of solubility units from mol % to volumetric rates

From Appendix A (also Example 3), for temperature of 30 °C (86 °F) and pressure of 7 MPa (1015 psia) the solubility of gases in mole % are:

Convert the solubility units from mol % to volumetric rates. Density of water is 1000 kg/m³ (62.4 lbm/ ft³) and MW = 18

 

 

SI Solution:

Assume a basis of one m³ of liquid water

 

Repeat the above calculation for each gas component and populate the second column from the right of Table 4.

 

Table 4 SI. Conversion of mol % solubility to volumetric

 

 

FPS Solution:

Assume a basis of one bbl of liquid water

 

Repeat the above calculation for each gas component and populate the second column from right of Table 4.

 

 

Table 4 FPS. Conversion of mol % solubility to volumetric

 

 

Summary

This TOTM presented basic thermodynamic models and estimation methods for the solubility of light hydrocarbon gases, i.e., methane through normal butane, in water as a function of pressure, temperature, and composition. By incorporating the composition dependent fugacity coefficients predicted by the SRK EOS, at high pressures the accuracy of results improved considerably. The tip demonstrated the use of solubility charts of Figures 2 and 3 (Table 2A) for compositions of Gas A to estimate solubility of another gas such as Gas B with different compositions.

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, and Applications in Gas Processing) courses.

By: Mahmood Moshfeghian, Ph.D.

 

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References:

1. Sloan, E.D., Koh, C.A., Clathrate Hydrates of Natural Gases, 3^rd ed., CRC Press, 2008.

2. Moshfeghian, M., https://petroskills.com/blog/entry/00_totm/jan21-fac-Solubility-of-Gases-in-WaterP1, Light Hydrocarbon gases, Jan 2021

3. ProMax 5.0, Build 5.0.20034.0, Bryan Research and Engineering, Inc., Bryan, Texas, 2020.

4. A. Danesh, PVT and phase behaviour of petroleum reservoir fluids. Elsevier, 1998.

 

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Appendix A

 

Figure 2. Variation of solubility of CH₄, C₂H₆, C₃H₈, iC₄H₁₀, and nC₄H₁₀ with pressure and at low temperatures

 

Figure 3. Variation of solubility of CH₄, C₂H₆, C₃H₈, iC₄H₁₀, and nC₄H₁₀ with pressure and at moderate temperatures

 

Table 1A. Fugacity coefficient for gas mixtures A and B at 30°C (86°F) predicted by the SRK EOS

 

Table 2A (SI). Solubility data for gas mixtures A and B calculated by ProMax [3]

 

Table 2A (FPS). Solubility data for gas mixtures A and B calculated by ProMax [3]