GPSA (Dehydration)

SECTION 20

Dehydration Natural gas and associated condensate are often produced from the reservoir saturated (in equilibrium) with water. In addition, the gas and condensate often contain CO 2  and H2S which might require removal. This is frequently accomplish accomplished ed with aqueous solutions such as amines, potassium carbonate, etc. which saturate the gas or condensate with water. Liquid hydrocarbons may also contain water downstream of product treaters or upon removal from underground storage.

Dehydration is the process used to remove water from natural gas and natural gas liquids (NGLs), and is required to:

• prevent formation formation of hydrates and condensation condensation of of free water in processing and transportat transportation ion facilities,

• meet a water content specifica specification, tion, and • prevent corrosion

FIG. 20-1 Nomenclature  A B C Cp Cg Cs Css CT D d

= = = = = = = = = =

EOS Fs G H

= = = =

area, m2 constant in Equation 20-14 constant in Equation 20-14 heat capacity, kJ/(kg • K) gravity correction factor for water content salinity correction factor for water content saturation correction factor for sieve temperature correction factor diameter, m depression of the water dewpoint or the gas hydrate freezing point, °C Equation of State  ————  —  — sizing parameter for packed towers, √ kg/(m • s) mass velocity, kg/(m2 • h) enthalpy, kJ/kg

Qv Qw Ss t T Trg v V • w W Wbbl Wr x X y z

= = = = = = = = = = = = = = = = γ = μ  = ρ =

∆H = latent heat of vaporization, kJ/kg K vs vapor/solid equilibrium K-value vs = L = length of packed bed, m Lg = glycol flow rate, m3/h LMTZ = length of packed bed mass transfer zone, m Ls = length of packed bed saturation zone, m • m = mass flow rate, kg/h MTZ = mass transfer zone M = molecular mass MI = molecular mass of inhibitor N = number of theoretical stages P = pressure, kPa (abs)

∆P = pressure drop, kPa q Q Qc Qhl Qr Qs Qsi

= = = = = = =

Qst = Qtr =

actual gas flow rate, m3/h heat duty, kW reflux condensing heat duty, kJ/m3 regeneration heat loss duty, kJ total regeneration heat duty, kJ/m3 sensible heat, kJ/m3 duty required to heat mole sieve to regeneration temperature, kJ duty required to heat vessel and piping to regeneration temperature, kJ total regeneration heat duty, kJ

vaporization of water heat duty, kJ/m3 desorption of water heat duty, kJ amount molecular sieve req’d in saturation zone, kg thickness of the vessel wall, mm temperat ure, °C temperature, regeneration temperature, °C vapor velocity, m/s superficial vapor velocity, m/h water rate, kg/h water content of gas, mg/Sm3 water content of gas, m3/Mm3 water removed per cycle, kg mole fraction in the liquid phase mass fraction in the liquid phase mole fraction in the gas phase compressibility factor specific gravity viscosity, mPa  • s density, kg/m3

Subscripts i o l v t CO2 H2S   HC s L R I H2O H rg f p i

20-1

= = = = = = = = = = = = = = = = = =

inlet outlet liquid vapor total carbon dioxide hydrogen sulfide hydrocarbon solid phase lean inhibitor rich inhibitor inhibitor water hydrate regeneration regeneration feed permeate any component in a mixture

 WATER CONTENT CONTENT OF GASES AND LIQUIDS

Techniques for dehydrating natural gas, associated gas condensate and NGLs include:

• Absorption using liquid desiccants,

 Water Solubility in Liquid Hydrocarbons

• Adsorption using solid desiccants, • Dehydration with CaCl2, • Dehydration by refrigeration refrigeration and inhibition, • Dehydration by membrane membrane permeation, • Dehydration by gas stripping, and • Dehydration by distillation.

Fig. 20-2 sh 20-2 shows ows the solubilities of water in sweet liquid hydrocarbons. It is is based on experimental data developed in GPA RR-169.1 In sour hydrocarbon hydrocarbon liquids, liquids, water solubility solubility can be substantially higher. GPA RR-622  provides water solubility data for selected hydrocarbons in both sweet and sour systems. Equationsof-State (EOS) may be used to estimate water solubilities in hydrocarbon systems. GPA RR-42 3  provides a comparison of experimental versus calculated values using a modified SoaveRedlich-Kwong (SRK) EOS. Results from EOS methods should

FIG. 20-2 Solubility of Water in Liquid Hydrocarbons

20-2

be used with caution and verified with experimental data when possible.

carbons, generalized pressure-temperatur pressur e-temperature e correlations are suitable for many applications. applications. Fig. 20-45 i  iss an example of one such correlation which has been widely used for many years in the design of “sweet” natural gas dehydrators. It is based on the work of McKetta and Wehe6 and Olds, et.al. 7 The gas gravity correlation should never be used to account for the presence of H2S and CO2 and may not always be adequate for certain hydrocarbon effects, especially for the prediction of water content at pressures above 10 000 kPa (abs). The hydrate formation line is approximate and should not be used to predict hydrate formation conditions.

Fig. 20-3 giv 20-3 gives es hydrocarbon solubilities in water, which in general are considerably considerably less than water in hydrocarbons hydrocarbons.. This figure is based on experimental data devel oped in GPA RR-169. 1 Some experimental data is available in GPA RR-62. Yaws, et. al.4 provide a general correlation which may be used to estimate the solubility of over 200 hydrocarbons in water.

 Water Content of Natural Gases The saturated water content of a gas depends on pressure, temperature, and composition. The effect of composition increases with pressure and is particularly important if the gas contains CO2  and/or H2S. For lean, sweet natural gases containing over 70% methane and small amounts of heavy hydro-

The   following example is used to illustrate the use o f Fig. The 20-4. 20-4.

Example 20-1 — Determine the saturated water content for a sweet lean hydrocarbon gas at 66°C and 6900 kPa (abs).

FIG. 20-3 Solubility of Hydrocarbons in Water

20-3

FIG. 20-4 Water Content of Hydrocarbon Gas

20-4

From Fig. 20-4, 20-4,

 Water Content of High CO2 /H2S Gases 3

W = 3520 mg/Sm

 Acid gas components components,, carbon dioxide (CO2) and hydrogen sulfide (H2S), increase the solubility of water in natural gas due to the attraction of water for these molecules. molecules. The equilibrium water content of an acid gas mixture varies significantly with pressure, temperature and mixture composition. Liquid CO 2 and H2S hold more water than gaseous CO 2 and H2S, but the opposite is true for hydrocarbons as shown by Kobayashi and K atz. atz.8 This eff ect ect is is seen in the t he several example systems shown in Figs. 20-5 th 20-5 throu rough gh 20-8.

For a 26 molecular mass gas, Cg = 0.98

(Fig. 20-4) 20-4)

W = (0.98)(3520) = 3450 mg/Sm3 For a gas in equilibrium with a 3% brine, Cs = 0.93

(Fig. 20-4) 20-4)

W = (0.93)(3520) = 3270 mg/Sm3

FIG. 20-5 Saturated Water Content of CO 2

100000

  s   a   g    t   e   w

CO2 Pure CO2, 18.3°C Pure CO2, 25°C Pure CO2, 31.1°C Pure CO2, 50°C Pure CO2, 73.8°C Methane

10000

73.8°C 50°C

   3

  m    S    /   r   e    t   a   w   g   m

31.1°C 25°C 1000

50°C 18.3°C

31.1°C 18.3°C

100 1000

10000 Pressure, kPa (abs)

100000

FIG. 20-6 Saturated Water Content of H 2S

1000000

  s   a   g    t   e   w

171°C 100000 138°C

   3

  m    S    /   r   e    t   a   w   g   m

104°C 71°C 10000 38°C

1000 1000

10000 Pressure, kPa (abs)

20-5

100000

Fig. 20-59,10,11  shows the water content in pure CO 2 (solid curves) at various temperatures and pressures. The water content of a light natural gas from Fig. 20-4 is shown for comparative purposes using dashed lines. At low pressure, the water content of CO2 decreases with increasing pressure as expected from ideal solubility. At higher pressures the water content in CO2  increases with increasing pressure due to the increased density of CO2 and the attraction of water for CO 2. The vertical dashed lines in Fig. 20-5 at 18.3°C and 25°C represent the change in water content due to the phase change from vapor to liquid. The critical temperature for CO2  is 31°C. Near the critical temperature and critical pressure, the density of CO 2

changes significantly with a small change in pressure resulting in a large change in water content. This ef fect is seen at 31°C and 50°C. Similar behavior is seen in Fig.  20-612 for H2S. Although 104.4°C is above the critical temperature for pure H 2S, the temperature is subcritical for the mixture of H2S and water as shown by Carroll and Mather. 13 In both Fig. 20-5 and 20-6, the curves were generated with the Yarrison Model (RR-200), 11 but modified as needed (particularly in the liquid state) to improve agreement with experimental data. Fig. 20-79,14,15 shows the saturated water contents of selected mixtures of CH4, CO2 and H2S versus pressure at 37.8°C and

FIG. 20-7 Experiment al Saturated Water Content of Mixtures at 37.8°C and 93.3°C 100000

5.5% CH4 + 0.3% C3 + 25% CO2 + 69.2%H2S

5.4% CH4 + 0.3% C3 + 49.5% CO2 + 44.8%H2S

93.3°C

  s   a   g    t   e   w

100% CO2

10000

9% CH4 + 10% CO2 + 81%H2S 30%CO2+26.3%H2S+C1

        3

  m    S    /   r   e    t   a   w   g   m

37.8°C

100% CH4 100% CO2 30% CH4 + 60% CO2 + 10%H2S

1000 90% CH4 + 10% CO2 100% CH4

100 1000

10000 Pressure, kPa (abs)

100000

FIG. 20-8 Saturated Water Content of CO 2 -Rich Mixtures at 93.3°C

30000

25000   s   a   g    t   e   w

20000

   3

  m    S    /   r   e    t   a   w   g   m

15000

Pure CO2 70% CO2 + 30% CH4

10000

10% CO2 + 90% CH4

5000

Pure CH4 0 0

5000

10000

15000

20000

25000

Pressures, kPa (abs)

20-6

30000

35000

40000

45000

93.3°C. Fig. 20-86,9,16 shows the water content of pure CH 4, CO2, and mixtures of CH 4 and CO2 at 93.3°C. Several significant observations can be made from these figures and other available data. Water content is not strongly affected by a low concentration of carbon dioxide in methane, but a low concentration of methane in CO2 can strongly affect the water content. Similar behavior is expected with H2S.

2. Corrections for H2S and CO2 should be applied when the gas mixture contains more than 5% H 2S and/or CO2 at pressures above 4800 kPa (abs). These corrections become increasingly significant at higher concentrations and higher pressures. 3. The addition of small amounts of CH 4 or N2  to CO2 or H2S can dramatically reduce the saturated water content compared to the pure acid gas.

1. Saturated water content of pure CO2 and H2S and mixtures containing high concentrations of acid gases can be significantly higher than that of sweet natural gas, particularly at pressures above about 4800 kPa (abs) at ambient temperatures.

 Acid gas water content is a very complex subject. The data and methods presented here should not be used for final design. Figs. 20-5, 20-6, 20-7 and 20-8 are all based on experimental data with some model predictions. A cursory study of

FIG. 20-9 Correlation for Estimating the Water Content of Acid Gas Mixtures

20-7

these figures reveals the complexities involved. Estimation of water content requires a careful study of the existing literature, availability of experimental data, and knowledge of the system phase behavior. Water content estimates for the condition/mixture of interest can be achieved through interpolation or extrapolation of data, simple correlations, or using equations of state. Interpolation or extrapolation requires due caution and careful treatment with understanding of the physical properties and phase behavior of the acid gas mixture. Additional experimental data is the best way to verify predicted values. Even the most sophisticated EOS techniques must be tuned to experimental data to accurately estimate water content.  An exhaustive list of references containing experimental water content data for natural gas mixtures has been compiled in reference 17. Numerous correlations have been proposed to estimate water content in acid gas mixtures (Robinson, et al., 18  Maddox, et al.,19 Carroll and Mather,13 Carroll,20 Wichert and Wichert, 5  Yarrison, et al.11 Several of these involve equations of state or hybrid approaches involving an equation of state for the gas and activity coefficient model for the liquid water phase. Other approaches provide graphs to enable quick approximate results. The method of Wichert and Wichert use s Fig. 20-9 to estimate water content of acid gas mixtures relative to water content in sweet gas. The method is applicable to mixtures containing both CO2 and H2S. With gases containing CO 2, the CO2 concentration is multiplied by 0.70 to obtain an “equivalent” H2S concentration as shown in Equation 20-1. yH2S (equiv) = (yCO2)(0.7) + (yH2S)

Eq 20-1

The method is limited to an H 2S equivalent of 50 mol% and applicable for temperatures from 10 to 177°C and pressures from 1400 to 69 000 kPa (abs). In comparison with 70 data points covering natural gas mixtures with up to 50% equivalent H2S concentrations from 38 to 107°C and 1400 to 41 000 kPa (abs), the error for the this method was randomly distributed with an average absolute error of 10% and a maximum of 37%.

Fig. 20-10 shows experimental water content data for various mixtures compared to the method in Fig. 20-9. Example 20-2 — Determine the saturated water content of gas containing 79% CH4, 12% CO2 & 9% H2S @ 49°C & 10 000 kPa (abs). First, the acid gas composition must be converted to an equivalent H2S using Equation 20-1. yH2S (equiv) = (yCO2)(0.7) + (yH2S) yH2S (equiv) = (0.7)(12) + 9 = 17.4% Enter the left side of Fig. 20-9 at 49°C and move horizontally to the %H2S equivalent line (17.4%). Proceed vertically to the 10 000 kPa (abs) pressure line and move left to the Water Content Ratio scale. The water content ratio is 1.15. Multiply this times the sweet gas water content at 49°C and 10 000 kPa (abs) from Fig. 20-4 (1265 kg/Sm3). Water content of acid gas-natural gas mixture = (1.15)(1265) = 1455 kg/Sm3

 Water Content in the Hydrate Region Fig. 20-4 is based on the assumption that the condensed water phase is a liquid. However, at temperatures below the hydrate temperature of the gas, the “condensed” phase will be a solid (hydrate). The water content of a gas in equilibrium with a hydrate will be lower than equilibrium with a metastable liquid. This is acknowledged in the “Warning” i n Fig. 20-4. Hydrate formation is a time dependent process. The rate at which hydrate crystals form depends upon several factors including gas composition, presence of crystal nucleation sites in the liquid phase, degree of agitation, etc. During this transient “hydrate formation period” the liquid water present is termed “metastable liquid”. Metastable water is liquid water which, at equilibrium, will exist as a hydrate. GPA RR-45,21 50,22  and 8023  present experimental data showing equilibrium water contents  of gases above hydrates. Data from GPA RR-50 is presented i n Fig. 20-11. For compara-

FIG. 20-10 Comparison of Experimental vs. Calculated Water Contents for Acid Gases  Water Content mg/Sm3 wet gas Mixture

T, °C

P, kPa

11% CO2/89% C1

37.8

13 789

652

690

666

11% CO2/89% C1

71.1

6 895

4 591

4 575

4 575

20% CO2/80% C1

37.8

13 789

652

738

716

20% CO2/80% C1

71.1

6 895

4 526

4 655

4 687

20% CO2/80% C1

71.1

13 789

2 761

3 114

3 082

8% H2S/92% C1

54.4

10 342

1 782

1 701

1 701

27.5% H2S/72.5% C1

71.1

9 597

3 965

4 173

4 173

17% H2S/83% C 1

71.1

6 964

4 687

4 703

4 783

C1/CO2/H2S 30%/60%/10%

37.8

7 584

1 300

1 380

1 252

C1/CO2/H2S 9%/10%/81%

37.8

13 100

7 095

NA

10 144

5.31% C 1/94.69% CO 2

25.0

10 342

1 753

NA

1 413

5.31% C 1/94.69% CO 2

50.0

13 789

2 643

NA

3 210

Experiment

20-8

 Wichert &  Wichert

 Yarrison, et al., (RR-200)

tive purposes, the metastable water content of a sweet gas from Fig. 20-4 is also shown. Water content of gases in the hydrate region is a strong function of composition. Fig. 20-11 should not be extrapolated to other compositions.

in gas and/or NGL systems can plug pipelines, equipment, and instruments, restricting or interrupting flow. There are three recognized crystalline structures for such hydrates. In both, water molecules build the lattice and hydrocarbons, nitrogen, CO2  and H2S occupy the cavities. Smaller molecules (CH4, C2H6, CO2, H2S) stabilize a body-centered cubic called Structure I. Larger molecules (C 3H8, i-C4H10, n-C4H10) form a diamond-lattice called Structure II.

When designing dehydration systems (particularly TEG systems) to meet extremely low water dewpoint specifications, it is necessary to determine the water content of the gas in equilibrium with a hydrate. If a metastable correlation is used, one will overestimate the saturated water content of the gas at the dewpoint specification. This, in turn, may result in a dehydration design which is unable to meet the required water removal. Where experimental data is unavailable, utilization of a sound thermodynamic-based correlation can provide an estimate of water content in equilibrium with hydrates.

Normal paraffin molecules larger than n-C4H10 do not form Structure I and II hydrates as they are too large to stabilize the lattice. However, some isoparaffins and cycloalkanes larger than pentane are known to form Structure H hydrates. 13 Gas composition determines structure type. Mixed gases will typically form Structure II. Limiting hydrate numbers (ratio of water molecules to molecules of included gaseous component) are calculated using the size of the gas molecules and the size of the cavities in H 2O lattice.

 Water Content Measurement Specifications for water content measurement are given in GPA Publication 2140. These include the Valve Freeze Method, the Bureau of Mines Dew Point Tester, and the Cobalt Bromide Method. Cobalt bromide color change occurs at about 25–30 mg/kg.

From a practical viewpoint, the structure type does not affect the appearance, properties, or problems caused by the hydrate. It does, however, have a significant effect on the pressure and temperature at which hydrates form. Structure II hydrates are more stable than Structure I. This is why gases containing C3H8 and i-C4H10 will form hydrates at higher temperatures than similar gas mixtures which do not contain these components. The effect of C 3H8  and i-C4H10  on hydrate formation conditions can be seen in Fig. 20-13. At 6900 kPa (abs), a 0.6 relative density gas (composition is shown in Fig. 20-16) has a hydrate formation temperature which is 7°C higher than pure methane.

There are several commercial instruments available for monitoring water content based on other principles. Measuring water contents of less than 20 ppmw or making dewpoint determinations at less than –40°C can be very difficult.

HYDRATES IN NATURAL GAS SYSTEMS  A hydrate is a physical combination of water and other small molecules to produce a solid which has an “ice-like” appearance but possesses a different structure than ice. Their formation

The presence of H 2S in natural gas mixtures results in a substantially warmer hydrate formation temperature at a

FIG. 20-11 Water Content of 5.31% C 3 /94.69% C1 Gas in Equilibrium with Hydrate

20-9

FIG. 20-12

FIG. 20-14

Conditions for Hydrate Formation for Light Gases

Permissible Expansion of a 0.6-Gravity Natural Gas Without Hydrate Formation

See Caution on Fig. 20-13

FIG. 20-13

FIG. 20-15

Pressure-Temperature Curves for Predicting Hydrate Formation

13

Permissible Expansion of a 0.7-Gravity Natural Gas Without Hydrate Formation

15

See Caution on Fig. 20-13

20-10

given pressure. CO 2, in general, has a much smaller impact and often reduces the hydrate formation temperature at fixed pressure for a hydrocarbon gas mixture.

FIG. 20 -16 Gas Compositions Used for Fig. 20-13 through 20-15

The conditions which affect hydrate formation are:

Mole Fraction C1

0.9267

0.8605

0.7350

• Gas or liquid must be at or below its water dew point or

C2

0.0529

0.0606

0.1340

saturation condition (NOTE: liquid water does not have to be present for hydrates to form)

C3

0.0138

0.0339

0.0690

iC4

0.0018

0.0084

0.0080

• Temperature

nC4

0.0034

0.0136

0.0240

• Pressure

nC5

0.0014

0.0230

0.0300

0.603

0.692

0.796

Primary Considerations

Rel. Den.

• Composition Secondary Considerations

Example 20-4  — The gas in Example 20-3 is to be expanded from 10 000 kPa (abs) to 3450 kPa (abs). What is the minimum initial temperature that will permit the expansion without hydrate formation?

• Mixing • Kinetics • Physical site for crystal formation and agglomeration such as a pipe elbow, orifice, thermowell, or line scale

• Salinity

The 10 000 kPa (abs) initial pressure line and the 3450 kPa (abs) final pressure line intersect just below the 45°C curve on Fig. 20-15.  Approximately 44°C is the minimum initial temperature.

In general, hydrate formation will occur as pressure increases and/or temperature decreases to the formation condition.

Example 20-5 — How far may a 0.6 relative density gas at 14 000 kPa (abs) and 40°C be expanded without hydrate formation?

Prediction of Sweet Natural Gas Hydrate Conditions

On Fig. 20-14 find the intersection of 14 000 initial pressure line with the 40°C initial temperature curve. Read on the x-axis the permissible final pressure of 7600 kPa (abs).

Fig. 20-12, based on experimental data, presents the hydrate pressure-temperature equilibrium curves for pure methane, ethane, propane, and for a nominal 70% ethane 30% propane mix. Fig. 20-13 through 20-15, based on gas gravity, may be used for first approximations of hydrate formation conditions and for estimating permissible expansion of sweet natural gases without the formation of hydrates. The conditions at which hydrates can form are strongly affected by gas composition. Compositions used for the construction of  Fig. 20-13 through Fig. 20-15 are given in Fig. 20-16. The gases are saturated with water.

Example 20-3 — Find the pressure at which hydrate forms for a gas with the following composition. T = 10°C. Mole Fraction 0.784 0.060 0.036 0.005 0.019 0.094 0.002

Component C1 C2 C3 iC4 nC4 N2 CO2 Total

Mole Mass 16.043 30.070 44.097 58.124 58.124 28.013 44.010

1.000

kg/kg-mol of Mixture 12.58 1.80 1.59 0.29 1.10 2.63 0.09 20.08

Example 20-6 — How far may a 0.6 relative density gas at 14 000 kPa (abs) and 60°C be expanded without hydrate formation? On Fig. 20-14, the 60°C initial temperature curve does not intersect the 14 000 kPa (abs) initial pressure line. Therefore, the gas may be expanded to atmospheric pressure without hydrate formation. Conditions predicted by Fig. 20-13 through 20-15 may be significantly in error for compositions other than those used to derive the charts. For more accurate determination of hydrate formation conditions, the following procedures should be followed. In addition, Fig. 20-14 and 20-15 do not account for liquid water and liquid hydrocarbons present or formed during the expansion. These can have a significant effect on the outlet temperature from the pressure reduction device.

Hydrate Prediction Based on Composition for Sweet Gases Several correlations have proven useful for predicting hydrate formation of sweet gases and gases containing minimal amounts of CO2 and/or H2S. The most reliable ones require a gas analysis. The Katz method 25,26 utilizes vapor solid equilibrium constants defined by the Equation 20-2. y K vs = x s

Mole mass (Mgas) of gas mixture = 20.08

γ =

Mgas Mair

=

20.08 28.964

= 0.693

From Fig. 20-13 at 10°C P = 2200 kPa (abs) for 0.7 relative density gas

Eq 20-2

WARNING: Not good for pure components — only mixtures.

The applicable K-value correlations for the hydrate forming molecules (methane, ethane, propane, isobutane, 27 normal butane,28 carbon dioxide, and hydrogen sulfide) are shown in Fig. 20-17 to 20-23. Normal butane cannot form a hydrate by itself but can contribute to hydrate formation in a mixture.

20-11

FIG. 20-17 Vapor-Solid Equilibrium Constants for Methane

FIG. 20-18 Vapor-Solid Equilibrium Constants for Ethane

20-12

FIG. 20 -19 Vapor-Solid Equilibrium Constants for Propane

20-13

FIG. 20-20

FIG. 20-22

Vapor-Solid Equilibrium Constants for Iso-Butane

Vapor-Solid Equilibrium Constants for Carbon Dioxide

FIG. 20-23 Vapor-Solid Equilibrium Constants for Hydrogen Sulfide

FIG. 20-21 Vapor-Solid Equilibrium Constants for N-Butane

20-14

For calculation purposes, all molecules too large to form hydrates have a K-value of infinity. These include all normal paraffin hydrocarbon molecules larger than normal butane. Nitrogen is assumed to be a non-hydrate former and is also assigned a K-value of infinity.

2. Assume some temperature and predict the hydrate formation pressure for this gas using the solid-vapor Kdata. Plot the results on Fig. 20-24. Sample calculations for 1380 and 2070 kPa (abs) are provided below. This calculation has been repeated for 2760, 3450, 5520 and 6890 kPa (abs) to develop Fig. 20-24.

The K vs values are used in a “dewpoint” equation to determine the hydrate temperature or pressure. The calculation is iterative and convergence is achieved when the following objective function (Equation 20-3) is satisfied.

1380 kPa (abs) T = 4°C

i=n

∑ (yi/K vs) = 1.0

Eq 20-3

i=1

Prudence should be exercised when some higher molecular weight isoparaffins and certain cycloalkanes are present since they can form structure H hydrates.

Methane Ethane Propane Isobutane n-Butane Nitrogen Carbon dioxide Total

Mole Fraction in Gas

2070 kPa (abs) Kvs

y/Kvs

Kvs

0.784 0.060 0.036 0.005 0.019 0.094 0.002

2.04 0.79 0.113 0.046 0.21 * 3.0

0.384 0.076 0.319 0.109 0.090 0.000 0.001

1.75 0.50 0.072 0.027 0.21 * 1.9

1.000

C1 C2 C3 iC4 nC4  C5 

0.9267 0.0529 0.0138 0.0018 0.0034 0.0014

Total

1.0000

2.25 0.50 0.055 0.0225

y/Kvs 0.4119 0.1058 0.2509 0.0800

Kvs 1.75 0.205 0.030 0.0105

0.8486

y/Kvs 0.5295 0.2580 0.4600 0.1714

1.4189

3. The intersection of the lines in Fig. 20-24 will be the point at which hydrates start to form. In this example, the result is 3450 kPa (abs) and 11°C.

2760 kPa (abs)

FIG. 20-24

y/Kvs

0.979

Kvs

Σy/K vs = 1.0 @ 1570 kPa (abs)

Example 20-7 — Calculate the pressure for hydrate formation at 10°C for a gas with the following composition. Component

y

2070 kPa (abs)

Solution Sketch for Example 20- 8

0.448 0.120 0.500 0.185 0.090 0.000 0.001 1.344

*Infinity Interpolating linearly, Σy/K vs = 1.0 at 2100 kPa (abs)

The experimentally observed hydrate-formation pressure at 10°C was 2240 kPa (abs).

Example 20-8  — The gas with the composition below is at 24 100 kPa (abs) and 66°C. What will be the hydrate conditions when this gas is expanded? Component

Mole Fraction

C1  C2  C3  iC4  nC4  nC5 

0.9267 0.0529 0.0138 0.0018 0.0034 0.0014

Total

1.0000

Fig. 20-14 would predict permissable expansion only to a pressure around 4800 kPa (abs). Note:

Solution Steps: 1. Make several adiabatic flash calculations at different pressures and plot on a pressure versus temperature graph. (See Fig. 20-24) Initial Pressure kPa (abs)

Initial Temperature °C

Final Pressure kPa (abs)

Final Temperature °C

24 100 24 100 24 100 24 100 24 100

66 66 66 66 66

2070 2760 3450 4140 4830

3 7 11 14 18

The Katz correlation is not recommended above 7000–  10 000 kPa (abs), depending on composition. Prediction of hydrate formation conditions at higher pressures requires the use of other methods. Sloan, et.al. 29 present an alternate set of Kvs values which, in general, are valid to 30 000 kPa (abs). McLeod & Campbell30 present experimental hydrate data for natural gas mixtures up to 70 000 kPa (abs) as well as a correlation for estimating high pressure hydrate formation conditions. Blanc & Tournier-Lasserve31  provide experimental hydrate data to 100 000 kPa (abs) and compare prediction correlations with experimental data.

20-15

Hydrate Predictions for High CO2 /H2S Content Gases

Example 20-9  — Estimate the hydrate formation temperature at 4200 kPa (abs) of a gas with the following analysis using Fig. 20-25.

The Katz method of predicting hydrate formation temperature gives reasonable results for sweet normal paraffin hydrocarbon gases. The Katz method should not be used for gases containing significant quantities of CO 2 and/or H2S despite the fact that K vs values are available for these components. Hydrate formation conditions for high CO2/H2S gases can vary significantly from those composed only of hydrocarbons. The addition of H 2S to a sweet natural gas mixture will generally increase the hydrate formation temperature at a fixed pressure. 32

Component N2  CO2  H2S C1  C2  C3  iC4  nC4  C5+ 

 A method by Baille & Wichert for predicting the temperature of high H2S content gases is shown in Fig. 20-25.33 This is based on the principle of adjusting the propane hydrate conditions to account for the presence of H 2S as illustrated in Example 20-9.

M = 19.75

FIG. 20-25 Hydrate Chart for Gases Containing H 2S

20-16

mol % 0.30 6.66 4.18 84.27 3.15 0.67 0.20 0.19 0.40

γ = 0.682

Solution Steps: 1. Enter left side of Fig. 20-25 at 4200 kPa (abs) and proceed to the H2S concentration line (4.18 mol%) 2. Proceed vertically to the relative density of the gas ( γ = 0.682) 3. Follow the diagonal guide line to the temperature at the bottom of the graph (T = 17.5°C). 4. Apply the C3 correction using the insert at the upper left. Enter the left hand side at the H 2S concentration and proceed to the C 3  concentration line (0.67%). Proceed down vertically to the system pressure and read the correction on the left hand scale (–1.5°C)

ture of a sweet natural gas. In this example, at 6900 kPa (abs), the addition of H 2S (10 mol%) to a sweet gas mixture increases the hydrate temperature by 8°C. On the other hand, CO 2 has a minor effect on the hydrate formation temperature and slightly decreases the hydrate temperature for both the “sweet” and “sour” gases in this case. EOS-based computer programs are probably the most consistent method of predicting hydrate formation temperatures today. Accuracy when compared to experimental data is usually ± 1°C. This is generally adequate for design.

Hydrate Inhibition The formation of hydrates can be prevented by:

The C3  temperature correction is negative when on the left hand side of the graph and positive on the right hand side. Note:

1. Maintaining the system temperature above the hydrate formation temperature by the use of a heater and/or insulation.

TH = 17.5 – 1.5 = 16°C Fig. 20-25 was developed based on calculated hydrate conditions using the Peng-Robinson EOS. It has proven quite accurate when compared to the limited amount of experimental data available. It should only be extrapolated beyond the experimental data base with caution. Fig. 20-2634 presents experimental hydrate formation data for three mixtures of methane, propane and hydrogen sulfide. Results of selected hydrate prediction methods are also shown. The addition of CO 2 to pure methane will slightly increase the hydrate temperature at a fixed pressure. 35 However, the addition of CO2 to a “typical” sweet natural gas mixture will often lower the hydrate formation temperature at a fixed pressure. Fig. 20-27 is provided to portray these compositional effects. The hydrate curves for four gas compositions are shown. These were generated using a commercial hydrate program employing the Peng-Robinson EOS. The four gas compositions are: Sweet Gas (0.6 rel. den. gas from Fig. 20-16) Sweet Gas containing 10% CO 2 Sour Gas containing 10% H2S Sour Gas containing 10% CO2 and 10% H2S Note that H 2S significantly increases the hydrate tempera-

2. Dehydrating the hydrocarbon fluid (gas and/or liquid) to eliminate the condensation of liquid or solid water. 3. Injection of a chemical inhibitor to prevent or mitigate hydrate formation. In some cases, heating or dehydration may not be practical or economically feasible. In these cases, chemical inhibition can be an effective method of preventing hydrate formation. Chemical inhibition utilizes injection of thermodynamic inhibitors (sometimes called equilibrium inhibitors) or low dosage hydrate inhibitors (LDHIs). Thermodynamic inhibitors are the traditional inhibitors (i.e., one of the glycols or methanol), which lower the temperature of hydrate formation. LDHIs are either kinetic hydrate inhibitors (KHIs) or antiagglomerants (AAs). They do not lower the temperature of hydrate formation, but do diminish its effect. KHIs lower the rate of hydrate formation, which inhibits its development for a defined duration. AAs allow the formation of hydrate crystals but restrict them to sub-millimeter size. Thermodynamic Inhibitors  — Inhibition utilizes injection of one of the glycols or methanol into a process stream where it can combine with the condensed aqueous phase to lower the hydrate formation temperature at a given pressure. Both

FIG. 20-26 Experimental vs. Predicted Hydrate Conditions for Gases Containing C 1, C 3, and H2S Experimental Data17

Composition, mol %

 Predicted Temperature, °C

γ

Temperature, °C

Pressure, kPa (abs)

Fig. 20-19

Equation 20-4

Fig. 20-31

4.174

0.649

4.6

706

NA

2.6

5.4

7.172

4.174

0.649

11

1419

5.0

8.4

11.3

88.654

7.172

4.174

0.649

14.2

2024

7.2

11.2

14.1

88.654

7.172

4.174

0.649

18

3367

11.7

14.9

18.4

81.009

7.016

11.975

0.696

10.4

817

1.1

5.1

10.8

81.009

7.016

11.975

0.696

19.5

2813

11.7

14.9

21.5

60.888

7.402

31.71

0.823

13.1

686

2.8

7.1

13.2

60.888

7.402

31.71

0.823

19.1

1445

8.3

15.3

20.3

60.888

7.402

31.71

0.823

24.3

2558

12.8

19.7

24.8

60.888

7.402

31.71

0.823

27.8

4275

16.7

24.1

28.7

C1

C3

88.654

7.172

88.654

H2S

20-17

glycol and methanol can be recovered with the aqueous phase, regenerated and reinjected. For continuous injection in services down to –40°C, one of the glycols usually offers an economic advantage versus methanol recovered by distillation. At cryogenic conditions (below –40°C) methanol usually is preferred because glycol’s viscosity makes effective separation difficult. Ethylene glycol (EG), diethylene glycol (DEG), and triethylene glycol (TEG) have been used for hydrate inhibition. The most popular has been ethylene glycol because of its lower cost, lower viscosity, and lower solubility in liquid hydrocarbons. Physical properties of methanol and methanol-water mixtures are given in Fig. 20-28 through Fig. 20-31. Physical properties of the most common glycols and glycol-water mixtures are given in Fig. 20-32 through Fig. 20-49. Tabular information for the pure glycols and methanol is provided in Fig. 20-50. Equilibrium inhibitors are used in both pipeline/flowline applications as well as in low temperature gas processing facilities. To be effective, the inhibitor must be present at the very point where the wet gas is cooled to its hydrate temperature. Fig. 20-51 shows a flow diagram for a typical EG injection system in a refrigeration plant. In these facilities, the glycol inhibitor is sprayed into the gas upstream of the exchanger. The exchanger type can be shell and tube, plate or printed circuit. As water condenses, the inhibitor is present to mix with the water and prevent hydrates. Injection must be in a manner to allow good distribution in the gas flow path. It is common practice to inject 2 to 3 times the glycol rate calculated from the correlations that follow. The viscosity of ethylene glycol and its aqueous solutions increases significantly as temperature decreases. This effect must be considered in the design and rating of exchangers in low temperature gas processing facilities.

The inhibitor and condensed water mixture is separated from the gas stream along with a separate liquid hydrocarbon stream. At this point, the water dew point of the gas stream is essentially equal to or slightly lower than the separation temperature. Glycol-water solutions and liquid hydrocarbons can emulsify when agitated or when expanded from a high pressure to a lower pressure, e.g., JT expansion valve. Careful separator design normally allows nearly complete recovery of the diluted glycol for regeneration and reinjection. The regenerator in a glycol injection system should be operated to produce a regenerated glycol solution that will have a freezing point below the minimum temper ature encountered in the system. This is typically 75–80 wt%. Fig. 20-52 shows the freezing point of various concentrations of glycol water solutions. The minimum inhibitor concentration in the free water phase may be approximated by Hammerschmidt’s equation. 36 d

=

XI =

K H XI MI (1 – XI)

Eq 20-4

dMI K H + dMI

Eq 20-5

Where K H for ethylene glycol and methanol = 1297. Earlier editions of the Engineering Data Book suggested a range of K H  values (1297–2222) for glycols. Higher values of K H result in lower concentrations of rich (diluted) glycol (X I in Equation 20-5) which, in turn, suggests a lower inhibitor injection rate. Experimental data suggests K H = 1297 is the correct

constant as illustrated in Fig. 20-53. In some eld operations,

FIG. 20-27 Hydrate Formation Conditions for Sweet Gas Showing Effects of CO 2 and H 2S

20-18

FIG. 20-28

FIG. 20-30

Density of Aqueous Methanol Solutions at Various Temperatures

Heat of Vaporization of Methanol Versus Temperature

FIG. 20-29

FIG. 20-31

Vapor Pressure of Aqueous Methanol Solutions at Various Temperatures

Freezing Points of Aqueous Methanol Solutions

20-19

Figures 20-32 through 20-43 are reproduced from Gas Conditioning Fact Book, 1962, and Figures 20-44 through 20-49 are reproduced from the Dow monoethylene, diethylene, and triethylene glycol guides, 2003, with permission from “The Dow Chemical Company” and subject to all warranty disclaimers therein.

FIG. 20-32

FIG. 20-34

Densities of Aqueous Ethylene Glycol Solutions

Densities of Aqueous Triethylene Glycol Solutions

FIG. 20-33

FIG. 20-35

Densities of Aqueous Diethylene Glycol Solutions

Viscosities of Aqueous Ethylene Glycol Solutions

20-20

FIG. 20-36

FIG. 20-38

Viscosities of Aqueous Diethylene Glycol Solutions

Heat Capacities of Aqueous Ethylene Glycol Solutions

FIG. 20-37

FIG. 20-39

Viscosities of Aqueous Triethylene Glycol Solutions

Heat Capacities of Aqueous Diethylene Glycol Solutions

20-21

FIG. 20-40

FIG. 20-42

Heat Capacities of Aqueous Triethylene Glycol Solutions

Thermal Conductivity of Diethylene Glycol–Water Mixtures

FIG. 20-43

FIG. 20-41

Thermal Conductivity of Triethylene Glycol–Water Mixtures

Thermal Conductivity of Ethylene Glycol–Water Mixtures

20-22

FIG. 20-44

FIG. 20-46

Vapor Pressures of Ethylene Glycol at Various Temperatures

Vapor Pressures of Aqueous Triethylene Glycol Solutions at Various Temperatures

FIG. 20-45

FIG. 20-47

Vapor Pressures of Aqueous Diethylene Glycol Solutions at Various Temperatures

Dew Points of Aqueous Ethylene Glycol Solutions at Various Contact Temperatures

20-23

FIG. 20-48

FIG. 20-49

Dew Points of Aqueous Diethylene Glycol Solutions at Various Contact Temperatures

Dew Points of Aqueous Triethylene Glycol Solutions at Various Contact Temperatures

FIG. 20-50 Physical Properties of Selected Glycols and Methanol Ethylene Glycol

Diethylene Glycol

Triethylene Glycol

Tetraethylene Glycol

Methanol

C2H6O2

C4H10O3

C6H14O4

C6H18O5

CH3OH

Molecular Mass

62.1

106.1

150.2

194.2

32.04

Boiling Point* at 760 mm Hg, °F

387.1

472.6

545.9

597.2

148.1

Boiling Point* at 760 mm Hg, °C

197.3

244.8

285.5

314

64.5

 Vapor Pressure at 77°F (25°C) mm Hg

0.12

<0.01

<0.01

<0.01

120

Density (g/cc) at 77°F (25°C) (g/cc) at 140°F (60°C)

1.110 1.085

1.113 1.088

1.119 1.092

1.120 1.092

0.790

kg/m3 at 77°F (25°C)

1110

1113

1119

1120

790

–13

–8

–7

–5.5

–97.8

—

–54

–58

–41

16.5 4.68

28.2 6.99

37.3 8.77

44.6 10.2

0.52

47

44

45

45

22.5

Refractive Index at 77°F (25°C)

1.430

1.446

1.454

1.457

0.328

Specific Heat at 77°F (25°C) kJ/(kg•K)

2.43

2.30

2.22

2.18

2.52

Flash Point, °C (PMCC)

116

124

177

204

12

Fire Point, °C (C.O.C.)

118

143

166

191

Formula

Freezing Point, °C Pour Point, °C  Viscosity in centipoise at 77°F (25°C) at 140°F (60°C) Surface Tension at 77°F (25°C), dynes/cm

Note: These properties are laboratory results on pure compounds or typical of the products, but should not be confused with, or regarded as, specifications. * Glycols decompose at temperatures below their atmospheric boiling point. Approximate decomposition temperatures are: Ethylene Glycol 165°C Diethylene Glycol 164°C Triethylene Glycol 207°C Tetraethylene Glycol 238°C

20-24

FIG. 20-51 Example Glycol Injection System

FIG. 20-52 Freezing Points of Aqueous Glycol Solutions however, hydrate formation has been prevented with glycol concentrations corresponding with K H values as high as 2222. This is because hydrate suppression with glycols depends on

the system’s physical and ow characteristics as well as gas and glycol properties. In addition, published experimental hydrate data is at the hydrate dissociation point not the point of hydrate formation. The dissociation temperature can be several degrees higher than the formation temperature depending on the system dynamics. It is recommended that a system be designed using a K H of 1297. Once the system is operating it may be possible to operate at lower injection rates. Equation 20-4 and Equation 20-5 should not be used beyond 20 wt% for methanol and 50 wt% for the glycols. For methanol concentrations up to about 50 wt%, the Nielsen-Bucklin equation37 provides better accuracy: d = – 72.0 ln (xH2O)

Eq 20-6

Note that “xH2O” in Equation 20-5 is a mole fraction, not a mass fraction. Fig. 20-54 shows the mol% vs. wt% relationship for both methanol and EG. Equations 20-7 and 20-8 may also be used.

mol% =

Wt% I M I

  

20-25

wt% I M I

+

(100 – wt% I) 18

(100)

 

Eq 20-7

wt% =

 

(mol% I) (M I)

 –4 to 19°C, and methanol concentration in the aqueous phase of about 3 to 73 mole percent. Data from similar temperatures (within 1°C) from various sources were combined and represented in Fig. 20-55 as a single average temperature. There is substantial scatter in the data sources, and the trends shown

(100) wt%

[(mol%  I) (M I) + 18 (100 – mol%  I)]

Eq 20-8 Where: M of methanol = 32 M of EG = 62

in the gure should be treated as approximate. The term on the

y-axis is a K-value (vapor-liquid equilibrium constant) with the units [(kg/106Sm3)/mol% MeOH in the aqueous phase]. The use of Fig. 20-55 is demonstrated in Example 20-10.

Fig. 20-54, Equations 20-7 and 20-8 are useful when using Equation 20-6, Figs. 20-53, 20-55 and 20-56.

Methanol losses to the hydrocarbon liquid phase ar e more difficult to predict. Solubility is a strong function of both the water phase and hydrocarbon phase compositions. Fig. 20-5642 presents experimental data 43,44,45,46,47,48 ,49 showing solubility of methanol in a various hydrocarbon liquids — paraffin and mixtures of paraffins and toluene. The y-axis is a distribution coefficient, mol fr. of methanol in the aqueous phase divided by the mol fr. of methanol in the hydrocarbon phase. Lower values for the distribution coefficient indicate higher hydrocarbon solubility. The use of a distribution coefficient to estimate methanol losses to the hydrocarbon liquid phase is shown in Example 20-10.

Maddox et al. 38 presents a method of estimating the required inhibitor concentration for both methanol and EG. The method is iterative but converges easily after a few iterations. Fig. 20-53 provides a comparison of various inhibitor correlations with experimental data. 39,40,41 Experimental data at very high inhibitor concentrations is limited. Once the required inhibitor concentration has been calculated, the mass of lean inhibitor solution required in the water phase may be calculated from Equation 20-8. mI =

XR • mH2O XL – XR

Solubility of EG in the liquid hydrocarbon phase is typically minimal.36 A solubility of 0.036 kg/m 3 of NGL is often used for design purposes. However, entrainment and other physical losses may result in total losses significantly higher than this.

Eq 20-9

The amount of inhibitor to be injected not only must be sufficient to prevent freezing of the inhibitor water phase, but also must be sufficient to provide for the equilibrium vapor phase content of the inhibitor and the solubility of the inhibitor in any liquid hydrocarbon. For methanol, the vapor pressure is

Example 20-10  — 2.83 (10 6) Sm3/day of natural gas leaves an offshore platform at 38°C and 8300 kPa (abs). The gas comes onshore at 4°C and 6200 kPa (abs). The hydrate temperature of the gas is 18°C.  Associated condensate production is 56 m 3/106 Sm3. The condensate has a density of 780 kg/m 3 a M of 140 and is substantially paraffinic. Calculate the amount of methanol and 80 wt% EG inhibitor required to prevent hydrate formation in the pipeline.

sufciently high that a signicant quantity of inhibitor will

be lost to the vapor phase. Methanol vaporization losses may be estimated from Fig. 20-55.42  The data in Fig. 20-55 is derived from a variety of sources,43,44,45,46,47,48 at pressures ranging from about 2000–34 000 kPa (abs), temperatures from about

FIG. 20-53 Hydrate Suppression vs. Inhibitor Concentration in Mol% 140 Hammerschmidt KH = 2222

120 Hammerschmidt KH = 1297

100 Nielsen & Bucklin

   C    °  ,    d

80 85 wt% MeOH RR-106

60 73.7 wt% MeOH RR-106

65 wt% MeOH RR-106

40 25 wt% MEG R -92 R

50 wt% MeOH RR-106

20 50 wt% MEG RR-92

50 wt% MeOH RR-92

25 wt% MeOH RR-92

0 0%

10%

20%

30%

40%

50%

60%

Inhibitor concentration, mol%

20-26

70%

80%

90%

d = 14°C, M = 32

Solution Steps:

Solving for XI,

Methanol

1. Calculate the amount of water condensed per day from Fig. 20-4, Win = 850 kg/10 6 Sm3 Wout = 150 kg/10 6 Sm3

From Equation 20-5, XI = 0.255, From Equation 20-6, mol fr. = 0.175 (use this value in subsequent calculations)

 ΔW = 700 kg/106 Sm3 Water condensed = (700)(2.83) = 1980 kg/d

From Fig. 20-54, wt% = 0.275

2. Calculate required methanol inhibitor concentration from Equation 20-5 and 20-6

3. Calculate mass rate of inhibitor in water phase from Equation 20-9 (assume 100% methanol is injected) mI =

FIG. 20-54 Weight % vs. Mol% for Methanol and EG Solutions

XL – XR

=

(0.275)(1980) (1 – 0.275)

=

750 kg/d

4. Estimate vaporization losses fro m Fig. 20-55. @ 4°C and 6200 kPa (ga), losses = (30 kg/10 6 Sm3)/ mol% MeOH in water phase

100%

daily losses = (30)(2.83)(17.5) = 1490 kg/d

90%

5. Estimate losses to hydrocarbon liquid phase from Fig. 2056.

80%

70%

@ 4°C and paraffinic fluid, Dist. Ratio = 110 mol % MeOH in hyd. liquid = 17.5/110 = 0.16 mol%

60%

     %     t     w

XR • mH2O

1 m3 of condensate has a mass of 780 kg

50%

= (780/140) =5.6 kmol/m 3 of condensate 40%

= (5.6)(0.0016) =0.009 kmol MeOH 30%

= (32)(0.009) = 0.29 kg/m 3

20%

Methanol

10%

Total MeOH losses to the hydrocarbon liquid phase = (0.29)(2.83)(56) = 46 kg MeOH/d

EG

Total methanol injection rate = 750 + 1490 + 46 = 2286 kg/d

0% 0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

mol %

FIG. 20-55 Ratio of Methanol Vapor Concentration to Methanol Liquid Concentration 200 15.3oC     )

   3

   d    i   u   q    i    l

  m  s    d   u   o    t    S  e   u   n   q   o   a    i    l    l   n    i    i    l   m   r   o   e   n   p   a    l    h   o   t   n   e   a   m    h    t    t   e   n   e   m  c   r   g   e    k   p   e    l   o    M

19.4oC

100

10.4oC

0.5 oC

19.4°C 15.3°C

    (

5.0 oC

10.4°C

  e   s   a    h    P   r   o   p   a    V   n    i    l   o   n   a    h    t   e    M

5.0°C 0.5°C

-4.3 oC

 -4.3°C Distributionof methanol between aqueous and vapor phase, from various sources including VLE and LLE data.

10 10

100 Pressure, bar

20-27

400

Methanol left in the gas phase can be recovered by condensation with the remaining water in a downstream chilling process. Likewise, the methanol in the condensate phase can be recovered by downstream water washing.

•

During regeneration, contaminants in the water phase (such as dissolved solids) leave with the water, not the methanol

•

Can be transported in the vapor phase (significant methanol vaporization at injection point) which is useful for removing hydrates that have formed downstream of the injection point in the system

80 wt% EG 1. Calculate required inhibitor concentration from Equation 20-5

Methanol Disadvantages

d = 14°C, M = 62, K H = 1297 Solving for XI, XI = 0.40 2. Calculate mass rate of inhibitor solution in water phase from Equation 20-9 mI =

(0.40)(1980)

(0.80 − 0.40)

= 1980 kg/d  Vaporization and liquid hydrocarbon losses are negligible. Hydrate inhibition with methanol or EG is widely used in pipelines as well as in gas processing plants. The choice of inhibitor is influenced by several factors. A few of these are listed below:

•

Higher inhibitor losses to the hydrocarbon vapor and liquid phases

•

More difficult to recover methanol from the aqueous phase

•

More toxic than EG

•

More flammable than EG (lower flash point)

Methanol losses to the hydrocarbon vapor and liquid phases has become a more significant issue due to increasingly stringent contaminant specifications for condensate, NGLs and natural gas. EG Advantages •

 Very low solubility losses to the hydrocarbon phases and generally not regarded as a contaminant

•

Much easier to recover from the water phase (regeneration)

•

Less toxic and less flammable than methanol

•

Can also provide corrosion inhibition for “top of the line corrosion” in pipelines

Methanol Advantages •

Generally less expensive than EG

•

Requires lower concentrations in the aqueous phase

•

Can inhibit to very low temperatures

•

Has a lower viscosity than EG

FIG. 20-56 Liquid-Liquid Methanol Distribution Ratios 1000

    )

  e   s   e   a   s    h   a    P   h   s   P   c   u   i   o   n   e   a   u   g   q   r    A   O   n   i   n    i

100

   H    H    O    O   e   e    M    M   n   n   o   i   o    i    t    t   c   c   a   a   r   f   r    f   e   l   e    l   o   o    M    M     (

  o    i    t   a    R   n   o    i    t   u    b    i   r    t   s    i    D

No Toluene 28-33mol% Toluene

10

50-70 mol% Toluene 70-80 mol% Toluene Distribution of methanol between aqueous and hydrocarbon phases, data from various sources. Hydrocarbon phases includes various alkane and cycloalkane compounds. Data shows the variation of distribution with changes in the amount of toluene in the hydrocarbon phase.

1 -50

-30

-10

10

Temperature (oC)

20-28

30

50

EG Disadvantages •

Higher concentrations required in the aqueous phase

•

Higher viscosity makes physical separation from hydrocarbon liquid phase more difficult

•

Is transported in the liquid phase so generally not very effective in removing hydrates that have formed downstream of the injection point in the system

•

During regeneration, contaminants in the water phase (such as dissolved solids) accumulate in the EG phase requiring special regeneration designs, typically vacuum systems

 As a general rule, EG is usually the first choice when continuous inhibition is required. Methanol is typically used when inhibition is only occasionally required, such as during periods of cold weather or during start-ups and shut-downs.

Inhibition with Electrolytes Hydrates are also inhibited by dissolved salts in the water phase. GPA RR-15650 presents experimental hydrate formation data for water solutions with various concentrations of NaCl, KCl and CaCl2. Electrolytes also provide hydrate inhibition in t he presence of equilibrium inhibitors such as methanol and EG. Fig. 20-57 shows the inhibition performance of four solutions of methanol and NaCl.

Low Dosage Hydrate Inhibitors (LDHIs) LDHIs can provide significant benefits compared to thermodynamic inhibitors including: •

Significantly lower inhibitor concentrations and therefore dosage rates. Concentrations range from 0.1 to 1.0 weight percent polymer in the free water phase, whereas alcohols can be as high as 50%.

•

Lower inhibitor loss caused by evaporation, particularly compared to methanol. LDHIs typically have a vapor pressure so low that no detectable amount of the LDHI is lost to the gas phase except via water droplet entrainment. Likewise, LDHIs have low solubility in liquid hydrocarbons. Most LDHIs partition less than 10% to the liquid HC phase. Most of the loss to the liquid HC phase is from water droplet entrainment or emulsification.

•

•

•

•

Increased production rates, where inhibitor injection capacity or flowline capacity is limited.

•

Lower toxicity of KHIs.

Kinetic Hydrate Inhibitors — KHIs were designed to inhibit hydrate formation in flowlines, pipelines, and downhole equipment operating within hydrate forming conditions such as subsea and cold weather environments. Their unique chemical structure significantly reduces the rate of nucleation and hydrate growth during conditions thermodynamically favorable for hydrate formation, without altering the thermodynamic hydrate formation conditions (i.e., temperature and pressure). This mechanism differs from methanol or glycol, which depress the thermodynamic hydrate formation temperature so that a flowline operates outside hydrate forming conditions. Some applications benefit from combining KHIs and methanol or glycol. The mixed inhibition package reduces the hydrate equilibrium temperature relative to uninhibited fluids, but does not reduce it below the operating temperature. Because there is some methanol or glycol, the lower hydrate equilibrium temperature results in a lower subcooling or driving force for hydrates than in a KHI only inhibition scheme. KHI is effective at the reduced subcooling in preventing the formation of hydrates in the flowing system. This is useful where the operating temperature is too low for KHI to function by itself and where methanol only strategies are not desirable, e.g., for reasons of limited storage volume or injection rates. KHIs Compared to Methanol or Glycols — KHIs inhibit hydrate formation at a concentration range of 0.1–1.0 weight percent polymer in the free water phase. For relative comparison, methanol or glycol typically may be required at concentrations ranging 20 to 50 weight percent respectively in the water phase. At the maximum recommended KHI dosage, the current inhibition capabilities are up to 16°C of subcooling in gas systems without acid gas components and less than 11°C in sour gas or oil systems. New KHIs continue to expand the region of effectiveness. Subcooling is the difference between the equilibrium hydrate temperature and the system temperature. The degree of subcooling that can be achieved with KHIs is time

FIG. 20-57 Hydrate Inhibition with Methanol and NaCI

11000

Reduced capital expenses through decreased chemical storage and injection rate requirements; and usually no need for regeneration because the chemicals are not currently recovered. However, in cases where the pipeline water is injected into a water disposal well system, Kinetic hydrate inhibitors (KHIs) in the aqueous phase may come out of solution if the reservoir temperature is above the KHI solution Cloud Point temperature (or Lower Critical Solution Temperature). If the KHI does come out of solution, it may cause a large decrease in the water injectivity due to plugging rock pores. In these cases, KHIs may have to be removed from the water before injection into the water disposal well system.

Pure Water

10000 20% NaCl

9000 20% MeOH

8000   a    P    k  ,   e   r   u   s   s   e   r    P

5% MeOH-15% NaCl

7000 15% MeOH-5% NaCl

6000 5000 4000 3000 2000 1000

These are especially appropriate for offshore where weight and space are critical to costs.

0 0

Reduced operating expenses in many cases through decreased chemical consumption and delivery frequency.

4

8

12

Temperature, °C

20-29

16

20

dependent and flow dependent. Shorter times allow higher subcooling values in flowing systems. Non-flowing Hold Times are often too short for pipelines to restart without intervention, such as, temporary glycol injection or lower pressure restart until the KHI has been re-distributed throughout the length of the pipeline. It is important that the design subcooling is consistent with the expected system residence time KHI Screening Considerations  — Although KHIs are applicable under most producing conditions, certain conditions must be considered when evaluating a potential application, which include water salinity, freezing conditions, hold time (i.e., period of effectiveness), water saturation, acid gas component concentrations and high temperature processes. •

KHIs are frequently antagonized by commonly used oilfield chemicals, such as, corrosion inhibitors, drilling fluid components, et cetera, particularly where these other chemicals are surface active. An extensive pre-screening program to determine the impact of KHIs on CIs, and vice versa is required. The testing can take six months to complete.

•

 At water salinity levels greater than approximately 17% NaCl, the polymer may come out of solution, thereby reducing KHI effectiveness.

•

 A solution of KHI in water does not provide protection from freezing or icing conditions, neither in the line being treated nor in the KHI storage tank. If ambient temperatures are expected to fall below freezing, the KHI storage volume must be freeze-protected through the use of insulation on the container and piping or addition of antifreeze (typically ethylene glycol) to the KHI solution.

•

•

 Antiagglomerant (AA) Inhibitors — Antiagglomerants were developed out of the necessity to extend the range of subcooling for LDHIs beyond that of KHIs, and AAs can achieve subcooling of greater than 4°C. Unlike KHIs, which delay the formation of hydrates, AAs allow their formation at normal rates, but as small nonagglomerating hydrate crystals that are dispersed into an oil or condensate preventing the formation and accumulation of large hydrate crystals. Thus, AAs are suitable only in the presence of liquid hydrocarbon or in nearly 100% water systems where less than 30% of the water is converted to hydrates due to limited quantities of gas. The mechanism of dispersion is emulsification with the AAs acting as emulsification agents.  AAs Compared to Methanol or Glycols — The comparisons of AAs to methanol and glycols are similar to KHIs except that AAs achieve greater subcooling and may not require thermodynamic inhibition after a shut-in/restart.  AA Screening Considerations — Although AAs are applicable under most producing conditions, certain conditions must be considered when evaluating a potential application. These conditions include water salinity, emulsification and demulsification (i.e., separation), pipeline hydraulics, water cuts, material compatibility, water treating, and downstream impacts.

 A solution of KHI cannot be used for melting ice or hydrate plugs. It is recommended to have other strategies, such as a sufficient quantity of ethylene glycol or methanol for remediation purposes in the event of a blockage. The KHI delivery system must be capable of providing sufficient dosage to achieve a hold time greater than the water residence time in the flowing piping. Factors to consider include: — The design basis for treating the pipeline with thermodynamic inhibitor or lower pressure during/after an unplanned shut-in.

•

Some AAs have a maximum or minimum salinity criterion. The maximum salinity criterion is normally not exceeded with produced water.

•

Since AAs are based on dispersing (i.d., emulsifying) polar hydrate crystals in a nonpolar oil or condensate phase (i.e., continuous phase), they may sometimes require a demulsifier for oil and water separation. Further, the addition of a heater upstream or heat coil inside a separator may be required to melt the hydrate crystals.

•

Since AAs form crystals that are then dispersed in the liquid hydrocarbon phase, careful consideration of the potential impact on viscosity should be considered including steady state flow, shut-in flow and restart conditions.

•

 An additional consideration for AAs is that the water cuts (i.e., percent water in the liquids) may have to be less than about 50% or higher than about 80% with limited gas fraction. Intermediate water cuts refer to concentrations below the inversion point (i.e.,where the continuous liquid phase changes from liquid hydrocarbon to water).  At the intermediate water cuts, AAs may fail to produce a flowable hydrate slurry for various reasons.

•

 AAs can impact the performance of some metallurgy and elastomers, so impacts on existing hardware should be reviewed.

•

Some oil-soluble AAs typically partition (i.e., disperse) to the liquid hydrocarbon phase, but low residuals can remain in the produced water, which can impact toxicity test results. Water-soluble AAs do present potential toxicity issues in the produced water.

•

Residual AA concentration in the hydrocarbon liquid phase could possibly impact downstream processes and should be considered in the context of overall contribution to a total feed-stream.

— The minimum required flow rate to entrain the KHI-treated liquids to the top-of-line. Liquid slugging can mitigate lack of droplet entrainment in gas systems. — The potential for water to pool in low sections of piping (e.g., turn-down hydraulics, flowline profile, pigging frequency, flowline interconnects that are not used continuously) and dead legs. — The seasonal duration of the cold point temperature below hydrate temperature, if applicable. •

•

If the gas is undersaturated with respect to water, the water in the KHI solution will evaporate and leave a high viscosity fluid. This can be addressed by using a more dilute KHI solution, or by changing the KHI carrier fluid to ethylene glycol. The KHI polymer and the solvents in the KHI product will form separate phases if the inhibited fluid is above the lower critical solution temperature (LCST) of the KHI/pipeline water solution.

The KHI polymer suffers degradation effects at temperatures above 249°C.

Chemical compatibility of LDHIs with other production chemicals can vary greatly; therefore, it is important to carry

20-30

out compatibility studies to investigate the effect of the LDHI on the performance of existing chemicals. The effect of the existing chemicals on the LDHI performance is equally important. The most frequently used production chemicals, especially in subsea pipelines, are corrosion inhibitors, scale inhibitors and paraffin inhibitors. These chemicals are either formulated in aqueous base solvents or hydrocarbon solvents.

GAS DEHYDRATION Glycol Dehydration Systems In those situations where inhibition is not feasible or practical, dehydration may be used. Both liquid and solid desiccants are employed, but economics frequently favor liquid desiccant dehydration when it will meet the required dehydration specification. Liquid desiccant dehydration equipment is simple to operate and maintain. It can easily be automated for unattended operation; e.g., glycol dehydration at a remote production facility. Liquid desiccants can be used for sour gases, but additional precautions in the design are needed due to the solubility of the acid gases in the desiccant solution. At high acid gas content and higher pressures the glycols can also be “soluble” in the gas. Glycols are typically used for applications where dew point depressions of the order of 33° to 66°C are required. Diethylene glycol (DEG), triethylene glycol (TEG), and tetraethylene glycol (TREG) are used as liquid desiccants, but TEG is the most common for natural gas dehydration. Following the process flow in   Fig. 20-58, the regenerated glycol (lean glycol) is pumped to the top of the contactor (absorber) which may contain trays or packing. The glycol absorbs water as it flows down through the contactor countercurrent to the gas flow. Water-rich glycol is removed from the bottom of the contactor, passes through the reflux condenser coil at the top of the regenerator, flashes off most of the soluble gas in the flash tank, and flows through the rich-lean heat exchanger to the regenerator. In the regenerator, absorbed water is distilled from the glycol at near atmospheric pressure by application of heat. The regenerated lean glycol exits the surge drum, is partly cooled in the lean-rich exchanger and is pumped through the glycol cooler before being re-circulated to the contactor.

refers to the difference between the actual outlet water dew point of the gas and the water dew point in equilibrium with the incoming lean glycol. Several equilibrium correlations 52,53,54,55,56,57,58 have been developed and published since 1950. All are limited by the ability to measure accurately the equilibrium concentration of water in the vapor phase above a highly concentrated TEG solution. Hence, accurate infinite dilution activity coefficients are critical to the reliable modeling of TEG dehydration systems. Parrish, et. al.51 and Bestani and Shing59 have reported infinite dilution activity coefficients for the TEG-water system. There is good agreement between these two sets of experimental data. The data of Bestani and Shing along with data from Herskowitz and Gottlieb were used to develop the activity coefficient correlation which provided the basis for the equilibrium calculations for Fig. 20-59. Equilibrium water dewpoints are relatively insensitive to pressure. Fig. 20-59 is based on a contactor pressure of 6900 kPa (abs). Higher pressures result in slightly lower dew points and lower pressures result in slightly higher dew points. For the pressure range 3450–13 800 kPa (abs) psia, the change in equilibrium dew points is typically less than 2–3°C. Please  note that the equilibrium water dew points on the ordinate of Fig. 20-59 are based on the assumption the condensed water phase is a metastable liquid. At low dew points the condensed phase will be a hydrate. The equilibrium dew point temperature above a hydrate is higher than that above a metastable liquid. Therefore, in these cases Fig. 20-59 predicts dew points which are lower than those which can actually be achieved. The difference is a function of temperature, pressure, and gas composition but can be as much as 8–11°C (se e Fig. 20-11). When dehydrating to very low dew points, such as those required upstream of a refrigeration process, the TEG concentration should be sufficient to dry the gas to the hydrate dew point not the metastable dew point. Once the lean TEG concentration has been established, the TEG circulation rate and number of trays (height of packing) must be determined. Most economical designs employ circula-

Good practice dictates installing an inlet gas separator upstream of the contactor, even if the dehydrator is near a production separator. The inlet gas separator will prevent carryover of liquid water (fresh or salty), hydrocarbons, treating chemicals or corrosion inhibitors into the glycol contactor. Even small quantities of these materials can result in excessive glycol losses due to foaming, reduced efficiency, and increased maintenance. Integral separators at the bottom of the contactor are common but it is important to ensure that the diameter and height of the separator are suitable for efficient liquid removal from the gas. Evaluation of a TEG system involves first establishing the minimum TEG concentration required to meet the outlet gas water dew point specification. Fig. 20-5951 shows the water dew point of a natural gas stream in equilibrium with a T EG solution at various temperatures and TEG concentrations. Fig. 2059 can be used to estimate the required TEG concentr ation for a particular application or the theoretical dew point depression for a given TEG concentration and contactor temperature. Actual outlet dew points depend on the TEG circulation rate and the number of equilibrium stages in the contactor, but typical approaches to equilibrium are 6–11°C. Approach to equilibrium

20-31

FIG. 20-58 Example Process Flow Diagram for Glycol Dehydration Unit

tion rates of about 15-30 kg lean TEG/kg H2O absorbed (13–27 liters lean TEG/ kg H 2O absorbed). The relationship between circulation rate and number of equilibrium stages is based on the absorption calculation techniques set out in Chapter 19. This has been done for TEG systems with the results presented in Figs. 20-60 through 20-6460.

The graphs in these figures apply only if the feed ga s is water saturated. The curves were developed based on contactor operating conditions of 6900 kPa (abs) and 40°C. The water removal is not a strong function of temperature and pressure, however a higher inlet gas water content results in a larger temperature rise across the contactor due to the heat of absorption of the water. This temperature rise decreases the absorption efficiency.

FIG. 20-59 Equilibrium H2O Dew Point vs. Temperature at Various TEG Concentrations 40.0 30.0 97.0 wt %

20.0

98.0 wt % 98.5wt %

10.0

99.0 wt %

0.0 99.5 wt %

   C    °-10.0  ,    t   n    i   o-20.0   p   w   e    D-30.0

99.7 wt %

99.9 wt % 99.95wt %

-40.0 -50.0

99.99 wt %

-60.0 -70.0 -80.0 10.0

20.0

30.0

40.0

50.0

60.0

70.0

Contactor Temperature, °C FIG. 20-60 Absorber Performance at Lean TEG Concentration = 98.5 w t% 1.00

0.90 N=3.0

  n   o    i    t   c   a   r    F0.80    l   a   v   o   m   e0.70    R   r   e    t   a    W 0.60

N=2.5 N=2.0

N=1.5

N=1.0 0.50 0

10

20

30

40

50

Circulation Ratio, mass lean TEG sol/mass water absorbed

20-32

60

70

In general, the correlations in Figs. 20-60 through 20-64 tend to overestimate water removal at lower feed gas p ressures and higher feed gas temperatures and underestimate water removal at higher pressures and lower temperatures. The variation depends on TEG concentration and circulation ratio but is typically less than 2-3%. Conversion from equilibrium stages to actual trays can be made assuming an overall tray efficiency of 25-30%. For packed columns it is more common to use transfer units than theoretical stages. Fig. 20-65 provides an approximate conversion from theoretical stages (N) to transfer units (NTU) as a function of TEG circulation ratio.

For TEG contactors, the height of a transfer unit (HTU) depends on the mass transfer rate. Increases in HTU represent a decrease in mass transfer. HTU decreases with increasing packing specific area and glycol circulation rate and increases with increasing gas rate and gas density. HTU is also affected by the TEG properties, particularly viscosity. Increases in TEG viscosity increase HTU. This may be a factor when operating at contactor temperatures less than 16°C. Fig. 20-66 can be used to estimate the HTU of structured packing in glycol contactors as a function of packing specific area and gas density.

FIG. 20-61 Absorber Performance at Lean TEG Concentration = 99.0 wt%

1.00 0.95 N=3.0

0.90   n   o    i    t   c0.85   a   r    F    l   a0.80   v   o   m   e0.75    R   r   e    t   a0.70    W

N=2.5 N=2.0

N=1.5

0.65 0.60 N=1.0 0.55 0

10

20

30

40

50

60

70

Circulation Ratio, mass lean TEG sol/mass water absorbed

FIG. 20-62 Absorber Performance at Lean TEG Concentration = 99.5 wt% 1.00

0.95

N=4.0

  n   o    i    t 0.90   c   a   r    F    l   a   v   o 0.85   m   e    R   r   e    t   a 0.80    W

N=3.0 N=2.5

N=2.0

0.75

N=1.5

0.70 0

10

20

30

40

50

Circulation Ratio, mass lean TEG sol/mass water abs orbed

20-33

60

70

The height of packing can be estimated from Equation 20-10.

It is common practice to add a minimum of two packing layers to the value calculated from Equation 20-10. The typical thickness of a layer of structured packing is 200 mm. For trayed contactors the typical tray spacing is 610 mm.

volume at the bottom of the column and, for packed columns, the liquid distributor. Bubble cap trays have historically been used in glycol contactors due to the low liquid rates versus gas flow, but structured packing is widely used today. Relative to bubble cap columns, structured packing typically allows a significantly smaller contactor diameter and a slightly smaller contactor height.

The total height of the contactor column will be based on the number of trays or packing required plus an additional 2–3 m to allow space for vapor disengagement above the top tray, inlet gas distribution below the bottom tray, rich glycol surge

Contactor diameter is set by the gas velocity. Sizing is identical to that outlined for separators in Section 7. Recommended values for K-factors and C-factors for glycol contactors are shown in Fig. 20-67.

Packing Height = (NTU)(HTU)

Eq. 20-10

FIG. 20-63 Absorber Performance at Lean TEG Concentration = 99.9 wt% 1.00

N=4.0

  n   o    i    t   c 0.95   a   r    F    l   a   v   o   m   e    R   r   e    t   a 0.90    W

N=3.5

N=3.0

N=2.5

N=2.0 0.85 0

10

20

30

40

50

60

70

Circulation Ratio, mass lean TEG sol/mass water absorbed

FIG. 20-64 Absorber Performance at Lean TEG Concentration = 99.99 wt % 1.00 0.99 0.98   n   o    i    t 0.97   c   a   r    F 0.96    l   a   v   o   m 0.95   e    R   r 0.94   e    t   a    W 0.93

N=4.0 N=3.5

N=3.0

0.92 N=2.5

0.91 0.90 0

10

20

30

40

50

Circulation Ratio, mass lean TEG sol/mass water absorbed

20-34

60

70

Structured packing vendors frequently quote an F s value for sizing glycol contactors, where Fs is defined in Equation 20-11. Fs = v (ρv)0.5

From Fig. 20-61 (99 wt% TEG) at a circulation ratio = 28 kg TEG/kg H2O, the required theoretical stages = 2.0 This is equivalent to 8 bubble cap trays. At 610 mm tray spacing this is 4.3 m between bottom and top tray.

Eq 20-11

 Values of Fs = 3.0 to 3.7 will generally provide a good estimate of contactor diameter for conventional structured packing. For high capacity structured packing, Fs values can range from 4.3 to 4.9. 6

For packing, from Fig. 20-65 at a circulation ratio = 28 kg TEG/kg H2O the required NTUs = 4.1 Feed gas density =

3

Example 20-11  — 0.85 10 Sm /d of a 0.65 sp gr natural gas enters a TEG contactor at 4100 kPa (abs) and 38°C. The outlet water content specification is 110 kg H2O/106 Sm3 and the TEG circulation rate is 28 kg lean TEG/ kg H 2O absorbed (25 liters lean TEG/kg H2O). Estimate the contactor diameter and number of bubble cap trays or height of structured packing required to meet this requirement. Assume z = 0.92 and that the specific area of the packing is 250 m2/m3.

  (4100)(0.65)(28.97) (0.92)(8.314)(273+38)

= 32 kg/m3

From Fig. 20-66, HTU = 0.76 m. Total height of packing = (4.1)(0.76) + 0.4 = 3.5 m The 0.4 m represents two additional 200 mm packing layers The total height is equivalent to 17-18 packing layers

Solutions Steps:

3. Estimate contactor diameter

1. Estimate required TEG concentration from Fig. 20-59

Bubble caps, 610 mm tray spacing:

H2O dew point = –4°C, which from Fig. 20-4 is equivalent to a water content of 110 mg H2O/Sm3 @ 4100 kPa (abs)

From Section 7 and Fig. 20-67:

Assume a 6°C approach to equilibrium so the equilibrium

G = C [ρv (ρL – ρv)]0.5

dew point is –10°C @ T = 38°C, lean TEG concentration ≈

= 176 [32 (1120 – 32)]0.5 = 32 800 kg/m2 • h

98.9 wt% (use 99.0 wt%) 2. Estimate number of theoretical stages.

·  = m

Calculate water removal efficiency From Fig. 20-4 the inlet gas water content at 38°C and 4100 kPa (abs) = 1450 kg/106 Sm3 Win – Wout Win

 1450–110 = 1450

 228 000 Sm3   1 kmol     23.64  d      (0.65)(28.97) kg  1 d      24 h  = 28 200 kg/h kmol   

= 0.924  A =

·  m G

=

28 200 32 800

= 0.86 m2

FIG. 20-65 Transfer Units (NTUs) vs. Theoretical Stages (N) 12.0

10.0

8.0

     s       U       T       N

6.0

N=4.0 N=3.5 N=3.0

4.0

N=2.5 N=2.0 2.0

N=1.5 N=1.0

0.0 0

10

20

30

40

50

Circulation Ratio, mass lean TEG sol/mass water absorbed

20-35

60

70

FIG. 20-66 Estimated HTU vs. Gas Density for Various Structured Packings

1.0

0.9

As = 250 m 2 /m 3 =(76 ft 2 /ft 3)

0.8   m  ,    U0.7    T    H

As = 300 m 2 /m 3 =(91 ft 2 /ft 3) As = 350 m 2 /m 3 =(107 ft 2 /ft 3)

0.6

0.5

0.4 0

25

50

75

100

125

Gas Density, kg/m3

 4 A  D =  π   

0.5

0.5

FIG. 20-67

(4)(0.86) = = 1.05 m   3.14   

Contactor Sizing Parameters for Glycol Contactors K factor, m/s

C factor, m/h

50 cm tray spacing

0.043

154

60 cm tray spacing

0.049

176

75 cm tray spacing

0.052

187

Structured (Standard)

0.091 to 0.122*

329–439

Structured (High Capacity)

0.122 to 0.152*

439-549

2.5 cm Pall rings

0.04 to 0.055

143–198

5 cm Pall rings

0.058 to 0.079

208–285

For Structured packing: Bubble Cap Trays

 CBubble cap 0.5 D =   (DBubble cap)  CStructured packing  =

 176 0.5   (1.05) = 0.7 m  384 

Packing 3

TEG will typically absorb about 3.7–4.5 Sm  of sweet natural gas per m3 of glycol at 6900 kPa (abs) and 38°C. Gas solubility will be considerably higher if the gas contains significant amounts of CO2 and H2S.

  Random

The estimated solubility of CO2 and H2S in TEG is shown in Figs 20-68 thru 20-71.61  These were developed using the data in GPA RR-183 and 189.62,63 Equation 20-1261 may also be used to estimate the solubility of H 2S and CO2 in TEG and EG as a function of pressure, temperature and acid gas composition in the vapor phase. At low acid gas partial pressures [less than 1400 kPa (abs)], the error is small, less than 10%. Accuracy decreases with increasing acid gas partial pressure.

* Depending on packing density and vendor

Pi = yi P yi is the mole fraction of acid gas in the vapor phase xi is the mole fraction of the acid gas in the liquid phase

(0.01) Pi xi =



exp  A + B





P 1 T – E  CxH2O + D T

T is the absolute temperature, K 



xH2O is the mole fraction of water in the liquid phase Eq 20-12

P is the absolute pressure, kPa (abs)

Pi is the partial pressure of the acid gas (component i: CO2 or H2S)

20-36

Glycol Reboiler Duty: Basis 1 m3 TEG

Constants for Equation 20-12 System H2S in EG/ Water H2S in TEG/ Water CO2 in EG/ Water CO2 in TEG/ Water

Sensible Heat:

A

B

C

D

E

4.256

–1298

1.824

0.02604

0.003354

Qs = m Cp ∆t = 3.037

–1414

1.4058

0.02130

0.003354

(1120 kg)  2.78 kJ   kg • °C  m3



(200ºC – 150ºC)



= 156 000 kJ/m3 6.125

–1232

1.287

0.01934

0.003354

4.813

–1239

1.429

0.01617

0.003354

Vaporization of Absorbed H2O: Qv = (ΔHvap)(ΔW) =

Flash tank sizing should be sufficient to degas the glycol solution and skim entrained liquid hydrocarbons, if necessary. A minimum retention time of 3–5 minutes is required for degassing. If liquid hydrocarbons are to be removed as well, retention times of 20–30 minutes may be required for adequate separation. Flash tank pressures are typically less than 520 kPa (abs).

 2260 kJ   kg H O  2  

 1 kg H2O  = 90 000 kJ/m3 0.025 m3 TEG  

Condenser Duty @ 25% Reflux Ratio: Qc = (0.25)(Qv) = 22 500 kg/m3 Total Duty Including 10% Heat Loss:

Regenerator sizing requires establishing the reboiler duty and, when high TEG concentrations are required, providing sufficient stripping gas.

Qr = (156 000 + 90 000 + 25 500)(1.1)

 A quick estimate of reboiler duty can be made using Equation 20-3.

Total Duty Based on 0.85 Sm3/d of gas:

Q = (418 000) (Lg )

= 295 000 kJ/m3

Q = (295 000 kJ/m3)(0.025 m3/kg) (0.85 106 Sm3/d/86 400) ((1450–110) kg/106 Sm3) = 97 kW

Eq 20-13

Equation 20-13 is approximate and usually gives values which are higher than the actual duty. A more rigorous determination of reboiler duty is shown in Example 20-12.

Example 20-12 — Determine reboiler duty for conditions in the previous example. Assume the rich TEG temperature entering the regenerator is 150°C and the reboiler temperature is 200°C.

Regeneration of TEG at 204°C and 1 atmosphere will produce about 98.6 wt% glycol. Regeneration at higher altitude (lower regenerator pressure) will result in higher concentrations at 204°C or a reduced regeneration temperature at the same concentration.

FIG. 20-68 Approximate Solubility of H 2S in Triethylene Glycol at 3500 kPa vs. Temperature, H 2S Content of Gas Phase, and Water Content of TEG

10.000

H 2 S 

  n   o    i    t   u    l   o   s    G 1.000    E    T

H 2 S 

H 2 S  H 2 S 

H 2 S 

H 2 S 

   3

  m    /    S    2    H    3   m 0.100  .    d    t   s  ,   y    t    i    l    i    b   u    l   o    S 0.010    S    2    H

H 2 S  H 2 S 

H 2 S 

H 2 S 

H 2 S  H 2 S 

0.001 20

30

40

50

60

Temperature,°C

20-37

70

80

90

100

CONTROL OF BTEX EMISSIONS FROM TEG REGENERATORS

 Vapor-liquid equilibrium constants (K-values) for benzene, toluene, ethylbenzene, and o-xylene (BTEX) in TEG solutions are presented in GPA RR-131.64

  Heavier paraffin hydrocarbons are essentially insoluble in TEG. Aromatic hydrocarbons, however, are very soluble in TEG, and significant amounts of aromatic hydrocarbons may be absorbed in the TEG at contactor conditions. This may present an environmental or safety hazard when they are discharged from the top of the regenerator.

Figs. 20-72 and 20-7365 show the estimated percentage absorption of BTEX  components in TEG at 2100 and 6900 kPa (abs) as a function of circulation ratio and contactor temperature. The circulation ratio is m 3/h of TEG solution/10 6 Sm3/d of gas. BTEX solubility increases with increasing circulation ratio

FIG. 20-69 Approximate Solubility of H 2S in Triethylene Glycol at 7000 kPa vs. Temperature, H 2S Content of Gas Phase, and Water Content of TEG

H 2 S 

  n10.00   o    i    t   u    l   o   s    G    E    T

H 2 S 

H 2 S 

   3

  m    / 1.00    S    2    H    3   m  .    d    t   s  ,   y    t    i    l    i    b   u    l   o 0.10    S    S    2    H

H 2 S 

H 2 S  H 2 S 

H 2 S  H 2 S 

H 2 S 

H 2 S 

H 2 S  H 2 S 

0.01 20

30

40

50

60

70

80

90

100

Temperature,°C

FIG. 20-70 Approximate Solubility of CO2 in Triethylene Glycol at 3500 kPa vs. Temperature, CO 2 Content of Gas Phase, and Water Content of TEG

C O  2 

  n   o    i    t   u    l   o   s 1.00    G    E    T

C O2  

C O2   C O  2  C O2  

   3

C O2  

  m    /    2

C O2  

   O    C    3

C O  2 

  m    d    t   s  ,   y 0.10    t    i    l    i    b   u    l   o    S

C O2  

C O2  

   2

   O    C

0.01 20

30

40

50

60

Temperature, °C

20-38

70

80

90

100

and decreasing temperature. The effect of pressure on solubility is more complicated. EOS correlations suggest solubility increases with increasing pressure at low pressures, but decreases with increasing pressure at pressures above 4800– 6900 kPa (abs). The two pressures used i n Figs 20-72 & 20-73 (2100 and 6900 kPa (abs)) are the two pressures at which experimental data was collected in GPA RR-131.

pendency is demonstrated with three sets of solubility curves (25°C, 35°C, and 50°C). In general, increased contactor temperature means higher inlet gas water content and a higher TEG circulation rate. This is why the 50°C curves are presented at high circulation ratios and the 25°C curves are presented at low circulation ratios.

Example 20-13 — Estimate the quantity of benzene absorbed in the glycol system presented in Example 20-11. The concentration of benzene in the feed gas is 300 ppm (0.0003 mol fr).

The BTEX absorption Figs 20-72 and 20-73 is presented for benzene, toluene, e-benzene and o-xylene. The temperature de-

FIG. 20-71 Approximate Solubility of CO2 in Triethylene Glycol at 7000 kPa vs. Temperature, CO 2 Content of Gas Phase, and Water Content of TEG 10.00

  n   o    i    t   u    l   o   s    G    E    T

C O  2  C O  2  C O  2  C O  2 

   3

  m    /    2

   O    C 1.00

C O  2 

   3

C O  2 

  m    d    t   s  ,   y    t    i    l    i    b   u    l   o    S

C O  2  C O  2  C O  2 

   2

   O    C

C O  2 

0.10 20

30

40

50

60

70

80

90

100

Temperature, °C

FIG. 20-72 BTEX Absorption in TEG Contactors at 300 psia ( 2068 kPa) Circulation Ratio, (m3 /h TEG)/(106 Sm 3 /d of gas) 0.0

0.4

0.8

1.2

1.6

2.0

2.4

60

50

60 Solid lines: 25°C (77°F) Dashed lines: 35°C (95°F) Dotted lines: 50°C (122°F)

50

   %  ,   n 40   o    i    t   p   r   o   s    b    A 30    X    E    T    B 20

   % 40  ,   n   o    i    t   p   r   o   s 30    b    A    X    E    T 20    B

10

0 0.00

2.8

10

0.05

0.10

0.15

0.20

0.25

Circulation Ratio, US GPM TEG/MMSCFD of gas

20-39

0.30

0 0.35

FIG. 20-73 BTEX Absorption in TEG Contactors at 1000 psia (6895 kPa) Circulation Ratio, (m3 /h TEG)/(106 Sm 3 /d of gas) 0.0

0.4

0.8

1.2

1.6

2.0

2.4

50

2.8 50

Solid lines: 25°C (77°F) Dashe d lines: 35°C (95°F) Dotte d lines: 50°C (122°F)

45

40

40

   %  ,   n   o    i    t 30   p   r   o   s    b    A    X    E20    T    B

35    %  ,   n   o    i    t 30   p   r   o   s 25    b    A    X 20    E    T    B 15

10

10 5

0 0.00

0.05

0.10

0.15

0.20

0.25

0.30

0 0.35

Circulation Ratio, US GPM TEG/MMSCFD of gas

Calculate TEG circulation rate in m3/h: (0.025 m3/kg H2O absorbed) • (0.85 106 Sm3/d)(1450–110) kg H2O/106 Sm3 = 1.2 m3/h 24 h/d

stream. Emission rates of BTEX pollutants increase rapidly with condenser operating temperature and stripping gas rate (if stripping gas is used). These factors will largely dictate the emission control performance for a given condenser and overhead vent stream.

Circulation ratio = (1.2 m3/h)/(0.85 106 Sm3/d) = 1.4

•

Air Cooled Condensers: Air cooled condensers with either natural or forced draft cooling are the most common options because of the relatively simple design and competitive capital cost. However, the condensing temperature will typically be 11 oC higher than the ambient dry bulb air temperature which limits the effectiveness of condensation in warm climates.

•

Glycol Cooled Condensers: Glycol cooled condensers use rich glycol from the contactor (prior to the glycol flash tank) as the coolant. The condenser outlet temperature is typically at or above the ambient temperature.

•

Water Quench Condensers: Water quench condensers combine a cool water stream with the regenerator overhead vent stream to condense the water and hydrocarbons. The quench water is cooled in a separate heat exchanger using air cooling or a refrigeration system.

•

Combined Air Cooled–Water Cooled Condenser: Combined air cooled–water cooled condensers use an air cooled condenser to partially condense the hydrocarbons and water. A water cooled condenser then cools the stream further and condenses essentially all the water and hydrocarbons present in the regenerator overhead vent stream. Cooling water is supplied to this condenser from a small cooling tower using water condensed from the regenerator overhead vent stream. The cooling water temperature will approach the wet bulb temperature which in dry climates will often allow the condensing temperature to be below the ambient air temperature.

From Fig. 20-72, estimated benzene absorption at 38°C and 2070 kPa (abs) is approx 9% From Fig. 20-73, estimated benzene absorption at 38°C and 6900 kPa (abs) is approx 10%  At 4100 kPa (abs) and 38°C use 10% absorption Mass flow of benzene in regenerator overhead stream. (850 000 Sm3/d)(0.0003)(78 kg/kmol)(10%) = 84 kg/d (23.64 Sm3/kmol) The feed gas to a glycol dehydration system will often contain small quantities of BTEX and as seen in Example 20-13, significant amounts of these BTEX components will be absorbed by glycol in the contactor. Almost all of the BTEX components are stripped from the TEG solution in the regenerator and will be present in the regenerator overhead vapor. In most countries, these components are considered hazardous air pollutants and emissions of these components are strictly regulated. Listed below is a summary of the commonly used emission control options. There are three main approaches to BTEX emission control from a glycol system. The first is to recover them as a salable liquid product by condensing the regenerator overhead vent stream. The second is to burn BTEX components in an incinerator/flare or as fuel. The third method is to recycle the components back into the process.

Condensers There are a variety of condensation systems using air, water, gas, or glycol to condense the regenerator overhead vent

If stripping gas is used in the regenerator there may be a significant amount of non-condensable hydrocarbons in the regenerator vent. This stream will be predominately methane

20-40

FIG. 20-74 Simplified Process Flow Diagrams of Enhanced TEG Regeneration Systems

but will also contain heavier hydrocarbons, including BTEX components. If the unit employs a direct-fired reboiler, the uncondensed vent stream is often used as fuel gas in the reboiler. If the reboiler heat source is hot oil, steam or electric coils, this vent stream is typically routed to a flare/incinerator or is recycled back into the process.

Recycle

Incinerators/Flares Regenerator overhead vent stream combustion is difficult because of the very high concentration of water in the stream and the possibility of hydrocarbons and water condensing in the incinerator/flare. Condensation in the incinerator/flare may result in smoking and incomplete combustion and in cold climates special precautions need to be taken to prevent freezing if the flare/incinerator is remote from the regenerator. Incinerators/ flares are used where it is uneconomical to recover the BTEX as a liquid product. Usually incinerators differ from flares in that they are designed to provide 99% + combustion of the flammable components in the feed. Catalytic combustion has been used in some locations.

In some facilities, a “near atmospheric pressure” vapor stream may exist in the process. A common example is the low pressure flash gas leaving the last stage of crude oil or condensate stabilization. A relatively simple method of handling the regenerator overhead vent stream is to use a blower or low pressure compressor such as a sliding vane or liquid ring compressor to boost the regenerator vapors into the flash gas steam. The hydrocarbons and water in the regenerator vent end up in the crude oil product or the produced water disposal system. The regenerator vent stream flow rate is typically small relative to the process stream.

ENHANCED GLYCOL CONCENTRATION PROCESSES There are several principles and processes for obtaining higher TEG purity than 98.7–99.0 wt%, which is the TEG purity obtained by regeneration at 204°C and atmospheric pressure.  All methods are based on the principle of reducing the effective

20-41

partial pressure of H2O in the vapor space of the glycol reboiler, and hence obtaining a higher glycol concentration at the same temperature.The most common method for enhancement of the glycol concentration is the use of stripping gas in the reboiler itself or in a packed stripping column below the reboiler. Other patented or proprietary processes in use to enhance the TEG purity and thereby achieve a lower water dew point are described below. The processes are illustrated on Fig. 20-74.

DRIZO

®

DRIZO® achieves glycol enrichment by means of a stripping medium which is a liquid at ambient conditions but a vapor at reboiler conditions. The stripping medium is a mixture of C 5+ hydrocarbons which may be supplied from external sources or internally generated by absorption into the TEG in the contactor and subsequent recovery as a liquid in the regenerator overhead. The condensed hydrocarbons, which contain BTEX components absorbed from the gas, are separated from the water in a 3-phase separator, vaporized and superheated, and flow to the lean-glycol stripping column to serve as the stripping medium. The advantage of the DRIZO ®  process is that very high stripping gas rates can be employed without additional hydrocarbon emissions from the regenerator. This can result in TEG purities above 99.99 wt %. In addition, the mixture of hydrocarbons and water form a heterogeneous azeotrope in the partial condenser at the top of the regenerator. As a result of this behavior, the condensing temperature is independent of the stripping rate. This is not the case when the stripping medium is natural gas.  As the hydrocarbons build up in the 3-phase separator, they are drawn off as a liquid stream which can be blended into a crude oil, condensate or NGL product. Various options to further enhance the lean TEG purity are available, such as drying the hydrocarbon liquid solvent with solid desiccant. In addition to the very high TEG concentrations DRIZO ® is effective in recovering BTEX components absorbed in the TEG.

tion or higher concentrations for a given stripping gas rate. Fig. 20-75 shows the approximate lean TEG concentration produced in an atmospheric pressure stripping column as a function of stripping gas rate. Lean TEG concentrations above 99.9 wt% have been achieved with stripping gas.

OTHER CONSIDERATIONS Under conventional dehydration conditions, 40 to 60% of methanol in the feed gas to a glycol dehydrator will be absorbed by the TEG. 66 This will add additional heat duty on the reboiler and additional vapor load on the regenerator. High methanol injection rates and slug carryover can cause flooding. Glycol losses can be defined as mechanical carryover from the contactor (normally 13 liters/106Sm3  for standard mist eliminator) plus vaporization from the contactor and regenerator and spillage. Glycol losses, exclusive of spillage, range from 7 liters/106Sm3 for high pressure low-temperature gases to as much as 40 liters/10 6Sm3 for low pressure, high temperature gases. Excessive losses usually result from foaming in the absorber and/or regenerator. Anti-foam agents are sometimes used. TEG vaporization losses at the contactor are minimal unless the gas temperature exceeds about 50°C. These losses are more significant at lower pressures. Tetraethylene glycol (TREG) has been used in some cases to minimize losses in high temperature, low pressure systems. Vaporization losses at the regenerator typically result from excessive stripping gas rates and/or inadequate reflux.

FIG. 20-75 Effect of Stripping Gas on TEG Concentration

COLDFINGER® The COLDFINGER ®  process can be described as a trace water exhauster for the TEG. A condensing tube bundle (“cold finger”) is inserted in the surge tank vapor space. In most applications, rich TEG from the glycol contactor is used as the coolant in the COLDFINGER ® tube bundle. The surge tank operates at reboiler temperature (193–204°C). The vapor in this space is water-rich, typically in excess of 50 wt% H 2O. Likewise, the liquid that condenses on the outside of the tube bundle has a relatively high H2O concentration. It is collected in a trough below the tube bundle and is removed from the surge tank, typically being recycled to the rich TEG regenerator feed stream. The liquid (lean TEG) in the bottom of the surge tank seeks to restore the equilibrium vapor composition and in doing so, exhausts its trace water before it reaches the surge-tank outlet nozzle. Operators of COLDFINGER ® units have reported lean TEG concentrations in the 99.2-99.5 wt% range. The COLDFINGER® process does not use stripping gas.

STRIPPING GAS Stripping gas is by far and away the commonly used technique for enhancing the lean TEG concentration. The stripping gas is typically dehydrated natural gas often taken from the fuel gas system. In rare cases, the stripping gas may be an inert gas, e.g. N2. Stripping gas can be introduced directly into the reboiler through a sparge tube in the lean TEG or it can be introduced in a countercurrent stripping column installed below the reboiler. The latter method provides more effective stripping which results in lower rates for a given lean TEG concentra-

20-42

Glycol losses in CO 2 dehydration systems can be significantly higher than in natural gas systems particularly at pressures above about 6200 kPa (abs). This is due to the solubility of TEG in dense phase CO 2. Glycerol67 is much less soluble in the vapor phase and has been used successfully as a desiccant in some CO2 dehydration systems. Glycol becomes corrosive with prolonged exposure to oxygen. A dry gas blanket on the glycol surge tank will help eliminate oxygen absorption. Special precautions should be taken if oxygen is in the gas to be dehydrated. Thermal decomposition of TEG can become a problem if TEG is heated to temperatures above 200°C.

dewpoint requirements, simultaneous control of water and hydrocarbon dewpoints, and special cases such as oxygen containing gases, etc. In processes where cryogenic temperatures are encountered, solid desiccant dehydration usually is preferred over conventional methanol injection to prevent hydrate and ice formation. Solid desiccants are also often used for the drying and sweetening of NGL liquids. Desiccants in common commercial use fall into one of three categories: Gels — alumina or silica gels manufactured and conditioned to have an affinity for water.  Alumina — a manufactured or natural occurring form of aluminum oxide that is activated by heating.

 A low pH accelerates decomposition of glycols. Bases such as triethanolamine, borax, or sodium mercaptobenzothiazole may be added to maintain pH, but they should be added sparingly.

Molecular Sieves — manufactured or naturally occurring aluminosilicates exhibiting a degree of selectivity based on crystalline structure in their adsorption of natural gas constituents.

SOLID DESICCANT DEHYDRATION There are several solid desiccants which possess the physical characteristic to adsorb water from natural gas. These desiccants generally are used in dehydration systems consisting of two or more towers and associated regeneration equipment. See Fig. 20-76 for a simple two-tower system. One tower is onstream adsorbing water from the gas while the other tower is being regenerated and cooled. Hot gas is used to drive off the adsorbed water from the desiccant, after which the tower is cooled with an unheated gas stream. The towers are switched before the onstream tower becomes water saturated. In this configuration, part of the dried gas is used for regeneration and cooling, and is recycled to the inlet separator. Other streams may be used if they are dry enough, such as part of the residue gas. Solid desiccant units generally cost more to buy and operate than glycol units. Therefore, their use is typically limited to applications such as high H 2S content gases, very low water

Silica Gel is a generic name for a gel manufactured from sulfuric acid and sodium silicate. It is essentially pure silicon dioxide, SiO2. It is used for gas and liquid dehydration and hydrocarbon (iC5+) recovery from natural gas. When used for hydrocarbon removal, the units are often called HRUs (Hydrocarbon Recovery Units) or SCUs (Short Cycle Units). When used for dehydration, silica gel will give outlet dewpoints of approximately –50°C.  Alumina is a hydrated form of alumina oxide (Al 2O3). It is used for gas and liquid dehydration and will give outlet dewpoints of about –68°C. Less heat is required to regenerate alumina and silica gel than for molecular sieve, and the regeneration temperature is lower. Molecular sieves give lower outlet water dewpoints. Molecular sieves are a class of aluminosilicates. They pro-

FIG. 20-76 Example Solid Desiccant Dehydrator Twin Tower System

20-43

duce the lowest water dewpoints and can be used to simultaneously sweeten and dry gases and liquids. Their equilibrium water capacity is much less dependent on adsorption temperature and relative humidity. They are usually more expensive. Molecular sieve dehydrators are commonly used upstream of deep recovery NGL extraction plants, typically cryogenic turbo-expander processes, as well as LNG liquefaction plants. These plants operate at very cold temperatures and require very dry feed gas to prevent formation of hydrates. Dehydration to a –100°C dewpoint is possible with molecular sieves. Water dewpoints less than –100°C can be accomplished with special design and strict operating parameters. Fig. 20-77 presents the important properties of commercial solid desiccants. Fig. 20-7868 shows static equilibrium capacity vs. relative humidity for various new desiccants. This is for comparison only. Adsorption capacity depends on other parameters besides relative humidity, and design values should be obtained from desiccant suppliers. Fig. 20-79 shows water-adsorption isotherms for 4A molecular sieve in contact with air. This gives the relationship of static-equilibrium water capacity to the operating temperature and gas dew-point temperature.  Although the chart is based on an air-water system, it also can be used for natural gas. The continuous process requires two (or more) vessels with one on-line removing water (in adsorption) while the other is being regenerated. Generally a bed is designed to be on-line in adsorption for 8 to 24 hours. When the bed is taken off-line, the water is removed by heating to 190°C–315°C, depending on the desiccant used and the performance specification (i.e., 190°C for silica gel and up to 315°C for molecular sieve, with alumina gel and activated alumina falling in between). The regeneration gas used to heat the bed is usually a slipstream of dry process gas. The regeneration gas is returned to the process after it has been cooled and the free water removed. Any heat source can be used to heat the regeneration gas including waste heat from engines and turbines, but red heaters are the most common heat source. This is an important design consideration since heat is often a major operating cost.

the desiccant will be at the top of the bed and will not affect the effluent dewpoint when adsorption is resumed. In addition, upflow heating helps to strip any contaminants from the top of the bed extending desiccant life. Regeneration gas flow during the cooling period may be upflow if the gas is completely free of water, which saves two switching valves per tower. If the cooling gas contains water, cooling flow should be downflow to avoid preloading of the desiccant at the bottom of the bed with water.

Design The first step is to determine the bed diameter, which depends on the superficial velocity. Too large a diameter will require a high regeneration gas rate to prevent channeling. Too small a diameter will cause too high a pressure drop and damage the sieve. The pressure drop is determined by a modified Ergun69  equation, which relates pressure drop to superficial velocity as follows:

∆P L

= B μ V + Cρ V2

Eq 20-14

Constants for Equation 20-11 are: Particle Type 3.2mm bead (4x8 mesh) 3.2mm extrudate 1.6mm bead (8x12 mesh) 1.6mm extrudate

B

C

0.0693 0.0893 0.1881 0.2945

3.75 x 10 –7 5.23 x 10 –7 5.74 x 10 –7 8.86 x 10 –7

Fig. 20-80 was derived from Equation 20-14 by assuming a gas composition and temperature and setting the maximum

allowable ∆P/L equal to 7.5 kPa/m. The design pressure drop

through the bed should be about 35 kPa. A design pressure drop higher than 55 kPa is not recommended as the desiccant is fragile and can be crushed by the total bed weight and pressure drop forces. Remember to check the pressure drop after the bed height has been determined. Once the allowable superficial velocity is estimated, calculate the bed minimum diameter (i.e., D minimum), and select the nearest standard diameter (i.e. D selected ):

Gas flow during adsorption is typically downflow. This allows higher gas velocities (thus smaller diameter towers) since bed fluidization is avoided. Regeneration gas flow is upflow during the heating period. In this way, any residual water left on

Dminimum = q =

·  m 60ρ

 4q  0.5 60 π Vmax  

Eq 20-15

Eq 20-16

FIG. 20-77 Typical Desiccant Properties

Heat Capacity, kJ/(kg  • K)

 Approx. Minimum Moisture Content of Effluent Gas (mg/kg)

Shape

Bulk Density, kg/m3

 Alumina Alcoa F200

Beads

770

7x14 Tyler mesh 3.2 mm/4.8 mm/6.4 mm

1.00

–68°C dew point

 Activated Alumina UOP A-201

Beads

735

3–6 mesh or 5–8 mesh

0.92

5–10 ppmv

Mole Sieve Grace – Davison 4A

Beads

675–720

4–8 mesh or 8–12 mesh

0.96

0.1 ppmv (–101°C)

Extrudate

640–705

3.2 mm/1.6 mm pellets

1.00

0.1 ppmv

Desiccant

Molecular Sieve UOP 4A-DG Mole Sieve Zeochem 4A

Particle Size

Beads

720–735

4–8 mesh or 8–12 mesh

1.00

0.1 ppmv

®

Beads

785

5x8 mesh

1.05

–51°C dew point

®

Silica Gel Sorbead  – H

Beads

720

5x8 mesh

1.05

–51°C dew point

Silica Gel Sorbead ® – WS

Beads

720

5x8 mesh

1.00

–51°C dew point

Silica Gel Sorbead  – R

20-44

FIG. 20-78 Static Equilibrium Capacity vs. Relative Humidity for Selected Solid Desiccants 46

Obtain the corresponding superficial velocity, V adjusted as follows:

 Dminimum 2 Vadjusted = Vmax  Dselected  

Eq 20-17



The next step is to choose an adsorption period and calculate the mass of desiccant required. Eight to twelve hour adsorption periods are common. Periods of greater than 12 hours may be  justified especially if the feed gas is not water saturated. Long adsorption periods mean fewer regeneration cycles and longer sieve life, but larger beds and additional capital investment. During the adsorption period, the bed can be thought of as operating with three zones. The top zone is called the saturation or equilibrium zone. The desiccant in this zone is in equilibrium with the wet inlet gas. The middle or mass transfer zone (MTZ) is where the water content of the gas is reduced from its inlet concentration to <1 ppmv. The bottom zone is unused desiccant and is often called the active zone. If the bed operates too long in adsorption, the mass transfer zone begins to move out the bottom of the bed causing a “breakthrough.” At breakthrough, the water content of the outlet gas begins to increase and will eventually reach feed gas water content when the MTZ is completely displaced. Both water capacity and the rate at which solid desiccants adsorb water decline as the material ages. The object of the design is to install enough desiccant such that after three to five years, the mass transfer zone will be at the bottom of the bed at the end of the adsorption period. In the saturation zone, molecular sieve is expected to hold approximately 13 kg of water per 100 kg of sieve. New sieve

FIG. 20-79 4A-DG MOLSIV

TM

UOP TM Adsorbents  Pellets — Water Adsorption Isotherms

20-45

will have an equilibrium capacity near 20%; 13% represents the approximate capacity of a 3–5 year old sieve. This capacity needs to be adjusted when th e gas is not water saturated or the temperature is above 24°C. See Fig. 20-81 and 20-82 to find the correction factors for molecular sieve. To determine the mass of desiccant required in the saturation zone, calculate the amount of water to be removed during the cycle and divide by the effective capacity. SS = LS =

Wr (0.13)(CSS)(CT )

π(D )(bulk density)

Even though the MTZ will contain some water (approximately 50% of the equilibrium capacity), the saturation zone is estimated assuming it will contain all the water to be removed. The length of the mass transfer zone can be estimated as follows: LMTZ = (Vadjusted/640 )0.3 (Z)

Eq 20-18 Where:

(Ss)(4) 2

Molecular sieve bulk density is 675–735 kg/m3 for spherical particles and 640–705 kg/m3 for extruded cylinders.

Eq 20-19

FIG. 20-80

Eq 20-20

Z = 0.52m for 3mm sieve 0.26m for 1.5mm sieve

The total bed height is the summation of the saturation zone and the mass transfer zone heights. It should be no less than the vessel inside diameter, or 1.8m, whichever is greater.

Allowable Velocity for Mole Sieve Dehydrator

Now the total bed pressure drop is checked. The ∆P/L for the selected diameter, D selected, is adjusted using Equation 20-14 or the following approximation:

 Vadjusted 2 (ΔP/L)adjusted ≅ (7.5 kPa/m)  Vmax  

Eq 20-21



The result is multiplied times the total bed height (L S + LMTZ) to get the total design pressure drop, which should be 35–  55 kPa. This is important, because the operating pressure drop can increase to as much as double the design value over three years. Too high a pressure drop plus the bed weight can crush the sieve. If the design pressure drop exceeds 55 kPa, the bed diameter should be increased and the sieve amount and vessel dimensions recalculated.  A second method uses Equation 20-17, but replaces the saturation capacity of 13% with an “effective desiccant capacity” which includes the MTZ effect, temperature, and relative humidity corrections. When using this method, an effective capacity of 8–10% is typically assumed. This method is adequate for most planning and feasibility calculations. To estimate the total cylindrical length of a tower, add a minimum of 1.0 m to the bed height, which provides the space for an inlet distributor and for bed support and hold-down balls under and on top of the sieve bed.

Regeneration Calculations FIG. 20-81 1 Mole Sieve Capacity Correction for Relative Water Saturation

The first step is to calculate the total heat required to desorb the water and heat the desiccant and vessel. A 10% heat loss is assumed.

FIG. 20-82 Mole Sieve Capacity Correction for Temperature

20-46

 kJ  4200 kg  (kg of water on bed)    1.0 kJ  Qsi = (kg of sieve)  (Trg – Ti) kg • K     0.5 kJ  Qst  = (kg of steel)  (Trg – Ti) kg • K    Qw

=

Qhl  = (heat loss) = Qw + Qsi = Qst)(0.10)

transfers to the bed, vessel steel, and heat loss to atmosphere; and the balance leaves with the hot gas.

Eq 20-22

The regeneration-gas flow rate is calculated from Equation 20-29 below. A typical average heat capacity for regeneration gas is 2.7 kJ/kg °C. The temperature, T hot, is 28°C above the temperature, Trg, to which the bed must be heated. The temperature, T b, is the bed temperature at the beginning of regeneration, which is the same as the dehydration-plant feed temperature.

Eq 20-23

Eq 20-24

Eq 20-25

The temperature, T rg, is the temperature to which the bed and vessel must be heated based on the vessel being externally insulated (i.e., no internal insulation which is usually the case). This is about 28°C below the temperature of the hot regeneration gas entering the tower. The weight of the vessel steel is estimated from the equations below. Equation 20-25 is the ASME Section VIII Div. 1 equation in terms of the vessel inside diameter. It is based on a maximum tensile stress of 130 MPa (i.e., the ASME 2001 maximum allowable tensile stress for SA516 Grade 70 steel at 343°C and a welded-joint efficiency of 1.0). This is 134 MPa at 315°C. The design pressure, P design, is usually set at 110% of the maximum operating pressure. The value of 3.2 mm in Equation 20-26 is the corrosion allowance in inches. The term 0.75D bed accounts for the weight of the tower heads. The value of 0.91 mm provides the space for the inlet distributor and support and hold-down balls. t(mm) = (1000 DbedPdesign) / ((2x130,000 = 260,000 kPa) – 1.2Pdesign)

Eq 20-26

Mass of steel (kg) = 29.8 (t + 3.2) (LS + LMTZ + 0.75 Dbed + 0.91)Dbed

Eq 20-27

• m rg = Qtr/(Cp(Thot – Tb) (heating time))

Eq 20-29

The superficial velocity of the regeneration gas is calculated from Equation 20-30 for which q is calculated from Equation 20-16. 4q V= Eq 20-30 (πD2) The calculated superficial velocity can not be less than the value that corresponds with a minimum bed pressure drop of 0.23 kPa/m. This can be determined from Fig. 20-84, which was derived from Equation 20-14 by assuming a gas composition

and temperature and setting ∆P/L equal to 0.23 kPa/m. If the

For determination of the regeneration gas rate, calculate the total regeneration load from Equation 20-28. Qtr = (2.5)(Qw + Qsi + Qst + Qhl)

The heating time is usually 50% to 60% of the total regeneration time which must include a cooling period. Fig. 20-83 shows a typical temperature profile for a regeneration period (heating and cooling). For 8 hour adsorption periods, the regeneration normally consists of 4 1/2 hours of heating, 3 hours of cooling and 1/2 hour for standby and switching. For longer periods the heating time can be lengthened as long as a minimum pressure drop of 0.23 kPa/m is maintained to ensure even flow distribution across the bed.

calculated velocity is less than this, the regeneration gas rate, ·  rg, must be increased by multiplying it by the ratio V min/V, and m the period of regeneration should be decreased by multiplying it times the ratio V/V min. Equation 20-14 may also be used to calculate Vmin using a ΔP/L equal to 0.23 kPa/m.

Eq 20-28

FIG. 20-84

The 2.5 factor corrects for the change in temperature difference (in – out) across the bed with time during the heating period. It assumes that 40% of the heat in the regeneration gas

Minimum Regeneration Velocity for Mole Sieve Dehydrator

FIG. 20-83 Inlet and Outlet Temperatures During Typical Solid Desiccant Bed Regeneration Period

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Vadjusted = (758 m/h)(2.21/2.25)2 = 731 m/h (Equation 20-17)

General Comments The regeneration cycle frequently includes depressuring/ repressuring to match the regeneration gas pressure and/or to maximize the regeneration gas volume to meet the velocity criterion. In these applications, the rate of depressuring or repressuring should not exceed 350 kPa per/minute. Some applications, termed pressure swing adsorption, regenerate the bed only with depressurization and sweeping the bed with gas just above atmospheric pressure, but this is not used in gas dehydration applications.

 

(∆P/L)adjusted = 7.5(731/758)2 = 7.0 kPa/m (Equation 20-21)

2. Estimate the amount of water to be removed from the feed per cycle for each bed. Base this on a 24-hour cycle consisting of 12 hours adsorbing and 12 hours regenerating (heating, cooling, standby, and valve switching). From Fig. 20-4, the water content at 4140 kPa (abs) and 38°C is 1410 mg/Sm3 (1410 kg/106 Sm 3). The water content at a dew point of –101°C is essentially zero, so the water removed is the following:

Moisture analyzers for very low water contents require care to prevent damage to the probes. When inserted into the beds, sample probes and temperature probes must be installed to reach the center of the gas phase.

w ˙ = (1410 kg/106 Sm3)(2.85 106 Sm3/day)/(24 h/day)/ = 167 kg/h of water removed

Solid desiccant towers are insulated externally or possibly internally. Internal refractory requires careful installation and curing, usually before the desiccant is installed. It saves energy but the greatest benefit is it can dramatically reduce the required heating and cooling times. This is often an important benefit for systems where regeneration times are limited. The primary disadvantage is the potential for wet gas bypassing the desiccant through cracks and defects in the insulation during the adsorption cycle.

Wr = (167 kg/h) (12 h) = 2004 kg water removed per 12-hour drying period or 24-hour cycle per bed. 3. Determine the amount of sieve required and the bed height based on a sieve bulk density of 720 kg/m 3. Since the feed gas is water saturated, the relative humidity is 100%, so CSS is 1.0 from Fig. 20-81. From Fig. 20-82, CT is 0.93 at 38°C. Applying the equations: SS = (2004)/((0.13)(1.0)(0.93)) = 16 576 kg of sieve for each bed (Equation 20-18)

Example 20-14: 2.85 × 106 Sm 3/day of natural gas with a molecular weight of 18 is to be processed for ethane recovery in a turbo-expander plant. It is water saturated at 4140 kPa (abs) and 38°C and must be dried to –101°C dew point. Determine the water content of the gas, and the amount of water that must be removed; and do a preliminary design of a molecular-sieve dehydration system consisting of two towers with down-flow adsorption in one tower and up-flow regeneration in the other. Use 4A molecular sieve of 3.2 mm beads (i.e., 4x8 mesh). The regeneration gas is part of the plant’s residue gas, which is at 4140 kPa (abs) and 38°C and has a molecular weight of 17. The bed must be heated to 260°C for regeneration.

LS = (16 576)(4)/(3.1416 (2.3)2 (720)) = 5.54 m bed height (Equation 20-19) LMTZ = (731/640)0.3 (0.52) = 0.54 m for mass-transfer zone (Equation 20-20) LS + LMTZ = 5.54 + 0.54 = 6.08 m of sieve for each bed The total sieve = (6.08/5.54)(16 576) = 18 192 kg for each bed

4. Check the bed design and pressure drop which is the ∆P/L calculated in Step 1 times the total bed height calculated in Step 3: (7.0 kPa/m)(6.08 m) = 42.6 kPa which meets the criteron of not exceeding 35–55 kPa

Solution Steps

1. Determine the bed diameter and the corresponding ∆P/L and V. First determine the maximum superficial velocity

5. Calculate the total heat required to desorb the water based on heating the bed and vessel to 260°C. First calculate the weight of steel from Equation 20-26 and 20-27. Let the design pressure, P design, be 110% of the operating pressure:

from Equation 20-14. Let the maximum ∆P/L be 7.5 kPa/ m. z = 0.93 from Fig. 23-8 for 17.4 mole weight which is conservative

Pdesign = (4140)(1.1) = 4554 kPa (ga)

ρ = (18 mole weight) (4140 kPa (abs))/((8.3145)

(311.15 °K) (0.93)) = 31 kg/m3 (Equation 23-2)

t=

μ = 0.015 mPa • s (Fig. 23-27)

(1000) (2.25) (4554) 260 000 – 1.2 (4554)

= 40.3 mm (Equation 20-26)

Mass of steel = (29.8) (40.3 + 3.2) (5.54 + 0.54 + (0.75) (2.25) + 0.91) (2.25) = 25 310 kg (Equation 20-27)

Vmax = [(7.5 kPa/m)/(3.75 • 10 –7)/(31 kg/m3)]1/2  – [(0.0693/3.75 • 10 –7) (0.015 mPa s)/ (31 kg/m3)/2] = 758 m/h (Equation 20-14), rewritten in terms of V

Qst = (25 310 kg) (0.5 kJ/kg °C) (260°C – 38°C) = 2 809 000 kJ (Equation 20-24)

m˙ = [(2.85  • 106 Sm3/day)/((24 h/day)(23.646

Sm3kmole))] (18 kg/kmole) = 90 400 kg/h

Qw = (4200 kJ/kg) (2004 kg water) = 8 417 000 kJ (Equation 20-22)

q = (90 400 kg/h)/(31 kg/m3) = 2916 m3/h (Equation 20-16)

Qsi = (18 192 kg) (1 kJ/kg °C) (260°C – 38°C) = 4 039 000 kJ (Equation 20-23)

Dminimum = [(4(2916 m3/h))/ (π • 758 m/h)]1/2 = 2.21 m (Equation 20-15)

Qhl = (2 809 000 + 8 417 000 + 4 039 000) (0.10) = 1 527 000 kJ (Equation 20-25)

Round off upward to 2.25 m diameter, for which V and ∆P/L are adjusted as follows:

Qtr = (2.5) (2 809 000 + 8 417 000 + 4 039 000 + 1 527 000) = 41 980 000 kJ (Equation 20-28)

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6. Calculate the flow rate of regeneration gas using Equation 20-29. Let the heating time be 60% of the total regeneration period.

Operating performance should be monitored periodically to adjust adsorption cycle length so that adequate dehydration is obtained. Performance tests are scheduled on a routine basis, ranging from monthly during early operations to six months or longer. The size of the unit and frequency of regeneration cycles also affect the timing of performance tests.

(60%) (12 hr) = 7.2 hours heating

m ˙ rg = (41 980 000 kJ)/((2.7 kJ/kg °C) (288°C – 38°C) (7.2 h)) = 8638 kg/h (Equation 20-29)

Refluxing Refluxing is a phenomenon that begins during the early heating stages in the regeneration of mol sieve adsorption beds. It is more common in systems where the regeneration is accomplished with high pressure gas, usually at or near feed gas pressure. At the beginning of the regeneration cycle, the hot gas picks up moisture and hydrocarbons as it flows upward through the bed. This gas cools as it travels upward through the bed and eventually reaches saturation at some point. Any cooling that takes place after saturation will result in the condensation of liquid water and hydrocarbons. This condensation is typically more prevalent at the vessel wall and on the vessel head at the top of the vessel. This temperature variance between the bottom of the bed and the top of the vessel results in significant condensation on the bed, at the side walls and the top section of the vessel.

7. Check that the ∆P/L ≥ 0.23 kPa/m at 228°C.  Assume z = 1.0.

ρ = ((17 mol weight)(4140 kPa (abs))/((8.3145) (273 + 288)(°K) • (1.0)) = 15.1 kg/m3

q = (8638 kg/h)/(15.1 kg/m3) = 572 m3/h of hot regeneration gas (Equation 20-16) Rearranging Equation 20-12: V = 4q/πD2 = ((4)(572)/((3.14)(2.25)2) = 144 m/h.

μ = 0.023 mPa • s (Fig. 23-27) ∆P/L = (0.0693) (0.023) (144) + (3.75 • 10 –7)

(15.1) (144)2 = 0.347 kPa/m (Equation 20-14)

This is safely above the minimum value of 0.23 kPa/m needed to prevent channeling. 8. The design results are summarized as follows: Number of vessels: two Vessel design pressure and temperature: 4554 kPa (ga) and 315°C Vessel dimensions: 2.25 m (2250 mm) ID by 6.99 m (6990 mm) tan to tan Weight of molecular sieve: 2 x 18,192 kg Regeneration gas rate: 8331 kg/h (2.78 • 105 Sm3/day) Regeneration gas temperature: 288°C Cycle time: 24 hours, 12 hours adsorption, 12 hours regeneration Heating time: 7.2 hours This example is based on the assumption that the regeneration gas is not recycled back to the feed gas. If the regeneration is recycled back the the feed gas, as illustrated on Fig. 20-76, the amount of regeneration gas must be added to the inlet gas. This is an iterative calculation. A good assumption is to assume that the regeneration gas rate is 10% of the feed rate. Bottom bed support typically includes three to five layers of inert ceramic balls in graduated sizes (smallest on top). On top of the bed, a hold-down screen is provided, again covered with a layer of ceramic balls. In some cases, a layer of less expensive desiccant can be installed on the top of the bed to catch contaminants such as free water, glycol, hydrocarbons, amines, etc. This may extend the bed life. Good inlet separation of entrained contaminants is absolutely essential for long desiccant life. Since solid desiccants can produce dust, 1 μm filters are frequently installed at the outlet of the dehydration unit to protect downstream equipment.

 As the regeneration cycle continues, the temperature of the vessel walls and the desiccant near the top of the bed increases. The results in re-vaporization of the condensed liquids, first the hydrocarbons and then the water. Certain components of the binders used in 4A sieve are somewhat soluble in liquid water. This solubility depends on the pH, temperature, and length of time that free water is in intimate contact with the adsorbent. Some of the soluble binder components can ion exchange with the zeolite and/or combine with anions in the liquid water to form solid when the water evaporates. This condition is made worse when contaminants such as amine or glycol are present due to carryover from an improperly designed inlet separator. These salts can cement the remaining molecular sieve pellets or beads together forming a solid mass that will restrict gas flow through the bed and reduce water capacity. Occasionally the solid salts form as a crust on the top of the bed or build-up along the vessel walls. Over time molecular sieve beds operating in this manner may exhibit broken particles, dust, increased pressure drop, channeling and premature breakthrough. Refluxing can be minimized by utilizing a two-step regeneration cycle where the bed is first heated to approximately 93°C with a lower temperature regeneration gas (approx. 121°C). Once the bed and vessel walls have reached the desired temperature, the regeneration gas temperature is increased to 288–  315°C and the heating cycle is completed as normal. This twostep process minimizes the opportunity for refluxing because there is a small temperature difference between the bottom and top of the bed when the bulk of the water desorption begins. Lower regeneration pressures are also helpful in preventing refluxing. These techniques increase the heating time required in the regeneration cycle and if the cycle time is limited they may be difficult to implement. Several mol sieve manufacturers have developed mol sieves that are more resistant to refluxing. As always, effective inlet gas separation is imperative for trouble-free operation of an adsorption dehydration system.

Calcium Chloride Calcium chloride (CaCl2) can be used as a consumable desiccant to dehydrate natural gas. Solid anhydrous CaCl2 combines with water to form various CaCl 2 hydrates (CaCl2 • XH2O). As

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water absorption continues, CaCl 2 is converted to successively higher states of hydration — eventually forming a CaCl2 brine solution. 10 to 20 mm CaCl 2 pellets are installed in a fixed bed much like a dry desiccant tower. Gas flow is upflow. The more effic ient designs utilize 3–4 trays below the solid bed to pre-contact the gas with the brine solution. This removes a portion of the water from the gas before contact with the sol id CaCl2 and increases unit capacity. One such unit is shown in  Fig. 20-85. The solid CaCl 2 near the bottom of the fixed bed will typically be CaCl2 • 4H2O or CaCl2 • 6H2O and the CaCl2  at the top of the fixed bed will be anhydrous CaCl2 or CaCl2 • H20. In this way the gas contacts successively drier CaCl 2  as it flows upwards and in theory leaves the fixed bed in equilibrium with the CaCl2 at the top of the bed. Outlet water contents of 16 mg/Sm 3  have been achieved with CaCl2 dehydrators. Typical CaCl2 capacity is 0.3 kg CaCl 2 per kg H2O. Superficial bed velocities are 6–9m and length to diameter ratio for the bed should be at least 3 to 4:1. CaCl2  dehydrators may offer a viable alternative to glycol units on low rate, remote dry gas wells. The CaCl 2  must be changed out periodically. In low capacity — high rate units this may be as often as every 2–3 weeks. Brine disposal raises environmental issues. In addition, under certain conditions the CaCl2 pellets can bond together to form a solid bridge in the fixed bed portion of the tower. This results in gas channeling and poor unit performance.

all the methanol in the cold decanted methanol water stream originating in the cold process at feed gas conditions to recirculate the methanol to the cold process. The water stream leaving the stripper contains generally less than 100 ppm wt of methanol. No heat is required for the process and no atmospheric venting takes place.

Dehydration By Membrane Permeation Membranes can be used to separate gas stream components in natural gas such as water, CO 2 and hydrocarbons according to their permeabilities. Each gas component entering the membrane unit has a characteristic permeation rate that is a function of its ability to dissolve in and diffuse through the membrane. The driving force for separation of a gas component in a mixture is the difference between its partial pressure across the membrane. As high pressure feed gas flows into the metal shell of the membrane unit, the fast components, such as water and CO2, permeate through the membrane. This permeate stream is collected at a reduced pressure, while the non-permeate stream, i.e., the dry natural gas, leaves the separator at a slightly lower pressure than the feed. The amount of co-permeation methane and other natural gas components is dependent on pressure difference between

FIG. 20-86  ® 

Example IFPEX-1  Dehydration Process Flow Diagram

Dehydration by Refrigeration The dehydration of natural gas can also be achieved by refrigeration and/or cryogenic processing down to –100°C in the presence of methanol hydrate and freeze protection. The condensed water and methanol streams decanted in the cold process can be regenerated by conventional distillation or by a patented process called IFPEX-1 ®. In the latter process illustrated in schematic form in  Fig. 208670 a slip stream of water saturated feed gas strips essentially

FIG. 20-85 Typical CaCl2 Dehydrator

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the feed gas and permeate stream, the surface area of the membrane and the membrane selectivity. For membrane systems designed to remove CO 2 from the gas, the co-permeation of hydrocarbons can be on the order of 10% of the feed gas stream. In CO2 units, almost all of the water will permeate the membrane with the CO2. For membrane units designed for dehydration only, the co-permeation of hydrocarbons is on the order of 1-2%, but a sweep gas stream is introduced into the permeate side of the membrane in order to remove water molecules from the membrane surface and thus increasing the permeation rate. The total of the co-permeation and sweep gas may be 4-5% of the feed stream. If there is no low pressure requirement for the permeate/ sweep gas stream then the gas must be flared or recycled using compression. This can significantly increase the cost and complexity of a membrane system.

LIQUID DEHYDRATION Many liquid streams must be dehydrated to allow further processing or meet requirements of a handling chain to a direct consumer. Commercial propane must be dry before entering the fuel market to prevent freezing problems as the liquid vaporizes at temperatures below the hydrate point, or even below the freezing point of any free water that may be present.

Special care must be taken in designing the bed supports in the liquid dehydrator vessels to prevent desiccant loss, desiccant damage, and to ensure proper distribution. Layers of ceramic balls are installed in decreasing size from the support screen. The support ball sizes may vary with the type and size of solid desiccant used but the layers of support balls should never be graduated in size more than twice the diameter of the balls being supported. The regeneration of solid desiccant beds is very similar to gas dehydrators with the following exceptions:

• Liquid draining and filling time must be allowed. • Pressuring and depressuring must be done carefully to avoid bed movement.

• Adequate bed cooling is required before liquid re-entry to minimize flashing. It is important to prevent movement of the bed particles to prevent attrition that would require premature replacement.  Also, desiccant dust particles can cause downstream plugging, equipment damage, and excessive filter maintenance. Liquid and vapor velocities must be controlled carefully and flashing of liquids or accelerated blow-down rates that would “lift” or “float” all or portions of the bed should be avoided. Desiccant bed life can be extended by doing several or all of the following activities:

• Prevent the desiccant particles from moving.

The amount of water that can be in solution with light hydrocarbon liquid is very small, even at the saturation point. Effective drying to very low levels of moisture is usually required. The solubility of water in liquid hydrocarbons is presented in Fig. 20-2. The desired maximum moisture level for commercial propane is  approximately 10 ppmw. However, liquids exposed to cryogenic temperatures require virtually all the moisture be removed.

•  Keep contaminants out of the dehydrating portion of the bed by upstream conditioning or by providing a sacrificial layer of less expensive desiccant to act as a catcher of any compounds such as amine, glycol or oil.

• Prevent overheating the bed to reduce the formation of carbon during the regeneration cycle.

• Analyze the heating/cooling regeneration tempera-

The water content in light hydrocarbon liquids can be determined by using recommended methods in GPA Publication 2140 (Cobalt Bromide or Freeze Valve methods), or electronic instruments designed to indicate the moisture content directly.

Gas Stripping One simple method of dehydrating liquid hydrocarbons is counter current stripping with a dry gas. This method is currently used to dry condensate produced offshore prior to export from the production platform. The contactor is usually trayed. Stripping gas rates depend on the condensate rate, the amount of entrained water in the condensate, stripper temperature and pressure.  Advantages of this process are simplicity and low capital cost. Disadvantages include the requirement for a dry natural gas stream, and the coincidental stripping of some of the volatile hydrocarbons from the condensate. The stripping gas may be recycled to the gas dehydration unit or it can be used as fuel gas.

Solid Desiccant Dehydration Several solid desiccant processes are available to dry liquid hydrocarbons. Liquid velocity is usually 1–1.5m/minute through solid desiccant beds with a minimum travel of at least 5 feet to ensure good distribution. Direction of flow can be upflow or downflow in the adsorption cycle.

ture cycles to minimize the time the bed is at elevated temperatures. This will also minimize energy requirements.  A typical heating/cooling regeneration temperature cycle plot is shown in Fig. 20-83. There are typically four (4) distinct stages in a normal cycle: Stage 1 — First bed-heating stage Stage 2 — Desorption stage Stage 3 — Second bed-heating stage Stage 4 — Bed-cooling stage For a period of time after the heat source is introduced into a desiccant bed being dehydrated, the bed must be heated to a temperature where the water will start to be desorbed (Stage 1).  As the water is desorbed (Stage 2), the bed temperature will usually rise only a few degrees because the regeneration gas heat is utilized to provide the heat of vaporization of the water being removed. The completion of the water desorption stage is characterized by a rapid increase in bed temperature measured as the outlet temperature. At this point the heating may be discontinued while bed heating will continue from residual heat in the heating cycle (Stage 3). As the unheated regeneration gas stream continues to pass through the bed, the bed will be cooled (Stage 4).  At near ambient pressures, regeneration of silica gel and alumina can be accomplished at 150°C. Molecular sieve requires 260–290°C to maintain the low dewpoint potential, and

20-51

the higher temperatures may increase desiccant life by providing more complete removal of adsorbed hydrocarbons. Capacity and performance data for new solid desiccants are usually presented based on a static test. Under operating conditions (dynamic) the performance data may be significantly different. Typically the effective capacity at operating conditions is about one-half of the capacity at equilibrium (static) conditions for most solid desiccants. This operating characteristic must be considered when designing a dehydration system and can be influenced by careful design and good control of operating parameters such as temperatures, contaminant levels, regeneration cycles, and desiccant selection. Solid desiccant manufacturers should be consulted for the most current product information and design criteria.

Distillation Wet NGLs can be dehydrated by distillation in specially designed fractionation columns. It will generally suffice to withdraw a sidestream liquid distillate three to four fractionation trays below the top of the distillation column to assure a dry product. Water in the NGL feed is passed overhead and is decanted in the overhead reflux drum while reflux is returned to the top of the column. Some extra costs will be required for the sidestream liquid withdrawal and cooling, but this can still be the most cost-effective manner to achieve the dehydration of propane and/or butane LPG products to adequate dryness specifications. The bottoms product of the distillation should be bone dry.

REFERENCES

Molecular Sieve

1. Wagner, Jan, RR-169, “Water-Hydrocarbon Mutual Solubility Data,” Gas Processors Association, June 1999.

Molecular sieve is not normally used for liquid dehydration because the required level of water removal is usually moderate and the cost of molecular sieve is considerably more than other types of suitable desiccants, such as activated alumina. However, in extreme cases where the moisture content of the liquid must be kept at an unusually low concentration, molecular sieve should be considered.

2. Brady, C. J., Cunningham, J. R., and Wilson, G. M., RR-62, “Water-Hydrocarbon Liquid-Liquid-Vapor Equilibrium Measurements to 530°F,” Gas Processors Association, Sept. 8, 1982. 3. Erbar, J. H., et.al., RR-42, “Predicting Synthetic Gas and Natural Gas Thermodynamic Properties Using a Modified Soave Redlich Kwong Equation of State,” Gas Processors Association, August 1980.

Molecular sieve may be used for removing other undesirable compounds, such as H 2S, COS, mercaptans, etc., from liquid streams. Dehydration may be a secondary benefit of using this type of treating method.

4. Yaws, C. L., et.al., “Hydrocarbons: Water Solubility Data,” Chem Eng. Vol. 97, No. 4, April 1990, p. 177. 5. Wichert, G. C., and Wichert, E., “New Charts Provide Accurate Estimations for Water Content of Sour Natural Gas,” O&GJ, October 27, 2003, pp 64–66.

Refer to the discussion of molecular sieve for gas dehydration elsewhere in this Section for more information.

 Activated Alumina

6. McKetta, J. J., and Wehe, A. H., “Use This Chart for Water Content of Natural Gases,” Petroleum Refiner (Hydrocarbon Processing), Vol. 37, No. 8, August 1958, p. 153.

There are several types of alumina available for use as a solid desiccant. Alumina is widely used for drying liquid product streams following gas processing, treating, or fractionation. Most alumina desiccants will produce a dewpoint below –70°C if applied properly. Alumina tends to adsorb heavy hydrocarbons which are difficult to remove during regeneration. Alumina is alkaline and is subject to reaction with mineral acids which are found in some well treating fluids.

7. Olds, R. H., Sage, B. H., and Lacy, W. N., “Composition of the Dew-Point Gas of the Methane-Water System,” Industrial and Engineering Chemistry, 1943, pp 1223–1227. 8. Kobayashi, R., and Katz, D. L., “Vapor-Liquid Equilibria for Binary Hydrocarbon-Water Systems,” I&EC, 45(2), 1953, pp 440–  451. 9. Wiebe, R., and Gaddy, V. L., “Vapor Phase Composition of Carbon Dioxide Water Mixtures at Various Temperatures and Pressures to 700 Atmospheres,” J. Am Chem Soc., Vol. 63, p. 475–477 (1941).

The design of a solid desiccant liquid dehydration system is similar to a gas dehydration system. An effective desiccant capacity of 4–5% is typically used in liquid dehydrator design.

Calcium Chloride

10. Gillespie, P. C., and Wilson, G. M., RR-41, “Vapor-Liquid Equilibrium Data on Water-Substitute Gas components: N 2-H2S, H2-H2O, CO-H2O, H2-CO-H2O, and H2S-H2O,” GPA Tulsa, OK 1980.

Calcium chloride is used as a consumable desiccant. Solid calcium chloride combines with water to form a brine solution. From one to two pounds of water can be absorbed by a pound of calcium chloride. Large liquid CaCl2  dehydrators are usually operated in a series that can be reversed with a moisture monitor located between the beds. In that way when the lead sacrificial bed is exhausted, no wet product is produced. The exhausted CaCl2 bed is then recharged and the vessels reversed in service.

11. Yarrison M., Song, K. Y., Cox, K., Chronister D., and Chapman, W., RR-200, “Water Content of High Pressure, High Temperature Methane, Ethane and Mathane+CO 2, Ethane + CO 2,” GPA, Tulsa, OK, March 2008. 12. Selleck, F. T., Carmichael, L. T., and Sage, B. H., “Phase Behavior in the Hydrogen Sulfide-Water System,” Ind & Engr. Chem.,  Vol. 44, No. 9, Sept. 1952, p. 2219.

The bed size can be estimated using a superficial velocity of 1 to 1.5m/min and an L/D of 2.5 to 4:1.

13. Carroll, J. J., and Mather, A. E. “Phase Equilibrium in the System Water-Hydrogen Sulfide: Modeling the Phase Behavior with an Equation of State,” Can. J. Chem. Eng. vol 67, 1989, pp 9991003.

Calcium chloride dehydration has become less popular because of the environmental problem of disposing of the produced brine.

14. Ng, H.-J., Chen, C,-J., and Schroeder, H., RR-174, “Water Content of Natural Gas Systems Containing Acid Gas, GPA, Tulsa, OK. 2001.

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15. Song, K. Y., and Kobayashi, R., “Water Content of CO 2 in Equilibrium with Liquid Water and/or Hydrates,” SPE Formation Evaluation, Dec. 1987, pp 500–508. 16. Kobayashi, R., and Song, K. Y., RR-120, “Water Content Values of a CO2 — 5.31 mol Percent Methane Mixture,” Gas Processors  Association, January 1989. 17. Kyoo Y. Song, Thi Vo, Matt Yarrison, Walter G. Chapman, “GPA PROJECT No. 975-6, Acid Gas Water Content, An Update of Engineering Data Book Information,” 2011. 18. Robinson, J. M., et.al, “Estimation of the Water Content of Sour Natural Gases,” Trans AIME, Vol. 263, August 1977, p. 281. 19. Maddox, R. N., et.al.; “Estimating Water Content of Sour Gas Mixtures,” Gas Conditioning Conference, Univ. of Oklahoma, Norman OK, March 1988. 20. Carroll, J. J., “The Water Content of Acid Gas and Sour Gas from 100 to 220 oF and Pressures to 10,000 psia,” presented at the 81st  Annual GPA Convention Dallas, Texas, March 11–13, 2002. 21. Aoyagi, et.al., RR-45, “I. The Water Content and Correlation of the Water Content of Methane in Equilibrium with Hydrates. II.The Water Content of a High Carbon Dioxide Simulated Prudhoe Bay Gas in Equilibrium with Hydrates,” Gas Processors Association, December 1980.

37. Nielsen, R. B., and Bucklin, R.W., “Why not use methanol for hydrate control?,” Hyd. Proc., Vol. 62, No. 4, April 1983, p. 71. 38. Maddox, R. N., et.al., “Predicting Hydrate Temperature at High Inhibitor Concentration,” Proceedings 1991 Gas Conditioning and Processing Conference, Univ. of Oklahoma, Norman, OK. 39. Ng, H.–J., Chen, C–J., Robinson, D. B., RR-92, “The Effect of Ethylene Glycol or Methanol or Hydrate Formation in Systems Containing Ethane, Propane, Carbon Dioxide, Hydrogen Sulfide or a Typical Gas Condensate,” Gas Processors Association, September 1985. 40. Ng. H.–J., Chen C.–J., Robinson, D. B., RR-106, “The Influence of High Concentrations of Methanol or Hydrate Formation and The Distribution of Glycol in Liquid-Liquid Mixtures,” Gas Processors Association, April 1987. 41. Défontaines, A. D., IFP Report No. 42324, “Experimental studies in hydrate formation in methanol hydrocarbons at very low temperatures,” July 1995. 42. Mamroush D. Trimeric, “GPA PROJECT No. 975-7, Methanol in Natural Gas, An Update of Engineering Data Book Information,”  A summary of data presented in RR-149 and RR-198, 2011. 43. Chapoy, A., et al., RR-198, “Water Inhibitor Distribution in Gas Production Systems,” Project 987, 2008.

22. Song, K. Y., and Kobayashi, R., RR-50, “Measurement and Interpretation of the Water Content of a Methane — 5.31 mol% Propane Mixture in the Gaseous State in Equilibrium with Hydrate,” Gas Processors Association, January 1982.

44. Ng, H.-J., et al., RR-66, “Equilibrium Phase Composition and Hydrating Conditions in Systems Containing Methanol, Light Hydrocarbons, Carbon Dioxide, and Hydrogen Sulfide,” Project 825, 1983 .

23. Song, K. Y., and Kobayashi, R., RR-80, “The Water Content of CO2  — Rich Fluids in Equilibrium with Liquid, Water, or Hydrate,” Gas Processors Association, May 1984.

45. Chen, C. J., et al., RR-117, “The Solubility of Methanol or Glycol in Water-Hydrocarbon Systems,” Project 825-87, 1988.

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46. Ng, H.-J., et al., RR-149, “Vapour-Liquid and Vapour-Liquid-Liquid Equilibria for H 2S, CO2, Selected Light Hydrocarbons and a Gas Condensate in Aqueous Methanol or Ethylene Glycol Solutions,” Project 905, 1995 . 47. Jacoby, R. H., “Vapor-Liquid Equilibrium Data for the Use of Methanol in Preventing Gas Hydrates,” Presented at the Gas Hydrate Control Conference. May 5–6, 1953. Norman, OK . 48. Nielsen, R. B., Aicher, T. G., Gray, R. M., “Methanol Hydrate Inhibition,” Presented at the Laurance Reid Gas Conditioning Conference, 2000, Norman, OK. 49. Bahadori, A., “Estimation of Hydrate Inhibitor Loss in Hydrocarbon Liquid Phase,” Petroleum Science and Technology. 27:9. 943–951. 50. Bishnoi, P. Raj, Dholabhai, Pankaj D., and Mahadev, Kal N., RR-156, “Hydrate Phase Equilibria in Inhibited and Brine Systems,” University of Calgary, Calgary, Alberta, Canada. Project 905-93 (Sub-Project 2). August 1996. 51. Clinton, P, Hubbard, R, and Shah, H, “A review of TEG-water equilibrium data and its effect on the design of glycol dehydration units,” Laurance Reid Gas Conditioning Conference, Norman, Oklahoma, Feb 2008.

33. Baillie, C., and Wichert, E., “Chart Gives Hydrate Formation Temperature for Natural Gas,” O&GJ, Vol. 85, No. 14, April 6, 1987, p. 37.

52. Parrish, W. R., Won, K. W., and Baltatu, M.E., “Phase Behavior of the Triethylene Glycol-Water System and Dehydration/Regeneration Design for Extremely Low Dew Point Requirements,” Proceedings 65th GPA Annual Convention, San Antonio, TX, 1986, p. 202.

34. Schroeter, et. al., TP-10, “Hydrate Decomposition Conditions in the System Hydrogen Sulfide, Methane and Propane,” Gas Processors Association, December 1982.

53. Townsend, F. M., “Vapor-Liquid Equilibrium Data for DEG and TEG-Water-Natural Gas System,” Proceedings 1953 Gas Conditioning Conference, Univ. of Oklahoma, Norman, OK.

35. Unruh, C. H., and Katz, D. L., “Gas Hydrates of Carbon DioxideMethane Mixtures,” Trans. AIME, Vol. 186, 1949, p. 83.

54. Scauzillo, F. R., “Equilibrium Ratios of Water in the Water-Triethylene Glycol-Natural Gas System,” Trans. AIME, Vol. 222, 1961, p. 697.

36. Hammerschmidt, E. G., “Formation of Gas Hydrates in Natural Gas Transmission Lines,” Ind. Eng. Chem., Vol. 26, 1934, p. 851.

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55. Worley, S., “Super Dehydration with Glycols,” Proceedings 1967 Gas Conditioning Conference, Univ. of Oklahoma, Norman, OK.