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 CO2 and H2S
which might require removal. This is frequently accomplished
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 of hydrates and condensation of free
water in processing and transportation facilities,
• meet a water content specification, and
• prevent corrosion
FIG. 20-1
Nomenclature
A
B
C
Cp
Cg
Cs
Css
CT
D
d
EOS
Fs
G
H
∆H
Kvs
L
Lg
LMTZ
Ls
•
m
MTZ
M
MI
N
P
∆P
q
Q
Qc
Qhl
Qr
Qs
Qsi
Qst
Qtr
= area, m
= 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)
2
= mass velocity, kg/(m • h)
= enthalpy, kJ/kg
= latent heat of vaporization, kJ/kg
= vapor/solid equilibrium K-value
= length of packed bed, m
= glycol flow rate, m3/h
= length of packed bed mass transfer zone, m
= length of packed bed saturation zone, m
= mass flow rate, kg/h
= mass transfer zone
= molecular mass
= molecular mass of inhibitor
= number of theoretical stages
= pressure, kPa (abs)
= pressure drop, kPa
= 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
Qv
Qw
Ss
t
T
Trg
v
V
•
w
W
Wbbl
Wr
x
X
y
z
γ
μ
ρ
2
= 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
= temperature, °C
= 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
feed
permeate
any component in a mixture
­ ATER CONTENT OF­
W
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 and inhibition,
• Dehydration by membrane permeation,
• Dehydration by gas stripping, and
• Dehydration by distillation.
Fig. 20-2 shows the solubilities of water in sweet liquid hydrocarbons. It is based on experimental data developed in GPA
RR-169.1 In sour hydrocarbon liquids, water 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-423 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
carbons, generalized pressure-temperature correlations are
suitable for many applications. Fig. 20-45 is 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.
be used with caution and verified with experimental data when
possible.
Fig. 20-3 gives hydrocarbon solubilities in water, which in
general are considerably less than water in hydrocarbons. This
figure is based on experimental data developed 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 of Fig.
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
­Water Content of High CO2/H2S Gases
­From Fig. 20-4,
­
W = 3520 mg/Sm
3
Acid gas 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. The equilibrium
water content of an acid gas mixture varies significantly with
pressure, temperature and mixture composition. Liquid CO2
and H2S hold more water than gaseous CO2 and H2S, but the
opposite is true for hydrocarbons as shown by Kobayashi and
Katz.8 This effect is seen in the several example systems shown
in Figs. 20-5 through 20-8.
­For a 26 molecular mass gas,
­
Cg = 0.98
(Fig. 20-4)
­
W = (0.98)(3520) = 3450 mg/Sm
3
­For a gas in equilibrium with a 3% brine,
Cs = 0.93
(Fig. 20-4)
W = (0.93)(3520) = 3270 mg/Sm
3
FIG. 20-5
Saturated Water Content of CO2
mg water / Sm 3 wet gas
100000
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
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 H2S
1000000
mg water / Sm3 wet gas
­
171° C
100000
138° C
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 CO2 (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 CO2. 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 CO2
changes significantly with a small change in pressure resulting
in a large change in water content. This effect 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 H2S,
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
Experimental 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
100% CO2
10000
9% CH4 + 10% CO2 + 81%H2S
30%CO2+26.3%H2S+C1
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 CO2-Rich Mixtures at 93.3°C
30000
25000
mg water / Sm3 wet gas
mg water / S m3 wet gas
93.3° C
20000
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 CH4, CO2,
and mixtures of CH4 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% H2S 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 CH4 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 uses 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 CO2, 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 H2S 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).
ater content of acid gas-natural gas mixture
W
= (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” in 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 in Fig. 20-11. For compara-
FIG. 20-10
Comparison of Experimental vs. Calculated Water Contents for Acid Gases
Mixture
T, °C
P, kPa
11% CO2/89% C1
Water Content mg/Sm3 wet gas
Experiment
Wichert &
Wichert
Yarrison, et al.,
(RR-200)
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% C1
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% C1/94.69% CO2
25.0
10 342
1 753
NA
1 413
50.0
13 789
2 643
NA
3 210
5.31% C1/94.69% CO2
20-8
in gas and/or NGL systems can plug pipelines, equipment, and
instruments, restricting or interrupting flow.
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.
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 (C3H8, 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 H2O 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 C3H8 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 H2S in natural gas mixtures results in
a substantially warmer hydrate formation temperature at a
FIG. 20-11
Water Content of 5.31% C3 /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. CO2, 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
Primary Considerations
•Gas or liquid must be at or below its water dew point or
saturation condition (NOTE: liquid water does not have
to be present for hydrates to form)
• Temperature
• Pressure
• Composition
Secondary Considerations
• Kinetics
•Physical site for crystal formation and agglomeration
such as a pipe elbow, orifice, thermowell, or line scale
• Salinity
In general, hydrate formation will occur as pressure increases and/or temperature decreases to the formation condition.
­ rediction of Sweet Natural ­
P
Gas Hydrate Conditions
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
Total
Mole
kg/kg-mol­
Mass of Mixture
16.043
12.58
30.070
1.80
44.097
1.59
58.124
0.29
58.124
1.10
28.013
2.63
44.010
0.09
1.000
Mgas
Mair
=
20.08
28.964
20.08
= 0.693
­From Fig. 20-13 at 10°C
­
0.8605
0.7350
C2
0.0529
0.0606
0.1340
C3
0.0138
0.0339
0.0690
iC4
0.0018
0.0084
0.0080
nC4
0.0034
0.0136
0.0240
nC5
0.0014
0.0230
0.0300
Rel. Den.
0.603
0.692
0.796
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.
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?
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).
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.
­ ydrate Prediction Based on ­
H
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 method25,26 utilizes vapor solid equilibrium constants defined by the Equation 20-2.
y
Kvs = x ­Mole mass (Mgas) of gas mixture = 20.08
γ =
0.9267
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
Component
C1
C2
C3
iC4
nC4
N2
CO2
C1
P = 2200 kPa (abs) for 0.7 relative density gas
s
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 Kvs 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.
i=n
∑ (yi/Kvs) = 1.0
1380 kPa (abs)
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
T = 4°C
y
Kvs
y/Kvs
Kvs
y/Kvs
C1
C2
C3
iC4
nC4
C5
0.9267
0.0529
0.0138
0.0018
0.0034
0.0014
2.25
0.50
0.055
0.0225
0.4119
0.1058
0.2509
0.0800
1.75
0.205
0.030
0.0105
0.5295
0.2580
0.4600
0.1714
Total
1.0000
0.8486
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.
Mole Fraction
in Gas
2070 kPa (abs) 2760 kPa (abs)
Kvs y/Kvs
Kvs
y/Kvs
FIG. 20-24
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
1.75
0.50
0.072
0.027
0.21
*
1.9
0.448
0.120
0.500
0.185
0.090
0.000
0.001
Solution Sketch for Example 20-8
1.000
0.384
0.076
0.319
0.109
0.090
0.000
0.001
0.979
1.4189
Σy/Kvs = 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
2070 kPa (abs)
1.344
*Infinity
Interpolating linearly, Σy/Kvs = 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
­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
Note: Fig. 20-14 would predict permissable expansion only to a
pressure around 4800 kPa (abs).
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
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 CO2 and/or H2S despite the
fact that Kvs 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 H2S to a sweet natural gas mixture will generally increase
the hydrate formation temperature at a fixed pressure.32
Example 20-9 — Estimate the hydrate formation temperature
at 4200 kPa (abs) of a gas with the following analysis using Fig.
­20-25.
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 H2S as illustrated in Example 20-9.
Component
mol %
N2
CO2
H2S
C1
C2
C3
iC4
nC4
C5+
0.30
6.66
4.18
84.27
3.15
0.67
0.20
0.19
0.40
M = 19.75
γ = 0.682
FIG. 20-25
Hydrate Chart for Gases Containing H2S
20-16
­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 H2S concentration and
proceed to the C3 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 H2S (10 mol%) to a sweet gas mixture increases
the hydrate temperature by 8°C. On the other hand, CO2 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:
Note: The C3 temperature correction is negative when on the
left hand side of the graph and positive on the right hand side.
­
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 con­
ditions 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 CO2 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% CO2
Sour Gas containing 10% H2S
Sour Gas containing 10% CO2 and 10% H2S
Note that H2S 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 C1, C3, and H2S
Experimental Data17
Composition, mol %
Temperature,­
°C
Predicted Temperature, °C
Pressure, ­
kPa (abs)
Fig. 20-19
4.6
706
NA
2.6
5.4
11
1419
5.0
8.4
11.3
2024
7.2
11.2
14.1
3367
11.7
14.9
18.4
1.1
5.1
10.8
11.7
14.9
21.5
2.8
7.1
13.2
C1
C3
H2S
γ
88.654
7.172
4.174
0.649
88.654
7.172
4.174
0.649
88.654
7.172
4.174
0.649
14.2
88.654
7.172
4.174
0.649
18
81.009
7.016
11.975
0.696
10.4
817
81.009
7.016
11.975
0.696
19.5
2813
60.888
7.402
31.71
0.823
13.1
686
Equation
20-4
Fig. 20-31
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
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 temperature 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 =
KH XI
MI (1 – XI)
dMI
XI =
K
+
dMI
H
Eq 20-4
Eq 20-5
Where K­H for ethylene glycol and methanol = 1297.
Earlier editions of the Engineering Data Book suggested a
range of KH values (1297–2222) for glycols. Higher values of
KH result in lower concentrations of rich (diluted) glycol (XI in
Equation 20-5) which, in turn, suggests a lower inhibitor injection rate. Experimental data suggests KH = 1297 is the correct
constant as illustrated in Fig. 20-53. In some field operations,
FIG. 20-27
Hydrate Formation Conditions for Sweet Gas Showing Effects of CO2 and H2S
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
Formula
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
Freezing Point, °C
–13
–8
–7
–5.5
–97.8
Pour Point, °C
—
–54
–58
–41
Viscosity in centipoise
at 77°F (25°C)
at 140°F (60°C)
16.5
4.68
28.2
6.99
37.3
8.77
44.6
10.2
0.52
Surface Tension at
77°F (25°C), dynes/cm
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
12
Flash Point, °C (PMCC)
116
124
177
204
Fire Point, °C (C.O.C.)
118
143
166
191
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 KH values as high as 2222.
This is because hydrate suppression with glycols depends on
the system’s physical and flow 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 KH 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% =


20-25
Wt% I
MI
wt% I
+
M I
(100)
(100 – wt% I) 
18

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
in the figure 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.
(100) wt%
[(mol% I) (M I) + 18 (100 – mol% I)]
Eq 20-8
Where: M of methanol = 32
M of EG = 62
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 are more difficult to predict. Solubility is a strong function of both the water
phase and hydrocarbon phase compositions. Fig. 20-5642 presents experimental data43,44,45,46,47,48 ,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/m3 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
sufficiently high that a significant 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
Example 20-10 — 2.83 (106) 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 m3/106
Sm3. The condensate has a density of 780 kg/m3 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.
FIG. 20-53
Hydrate Suppression vs. Inhibitor Concentration in Mol%
140
Hammerschmidt
KH = 2222
120
Hammerschmidt
KH = 1297
100
d, °C
Nielsen & Bucklin
80
85 wt% MeOH RR-106
60
73.7 wt% MeOH RR-106
65 wt% MeOH RR-106
40
25 wt% ME G R -R92
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/106 Sm3
Wout = 150 kg/106 Sm3
ΔW = 700 kg/106 Sm3
Water condensed = (700)(2.83) = 1980 kg/d
From Equation 20-5, XI = 0.255,
From Equation 20-6, mol fr. = 0.175 (use this value in
subsequent calculations)
From Fig. 20-54, wt% = 0.275
3. Calculate mass rate of inhibitor in water phase from
Equation 20-9 (assume 100% methanol is injected)
2. Calculate required methanol inhibitor concentration
from Equation 20-5 and 20-6
XR • mH2O
(0.275)(1980)
=
=
= 750 kg/d
mI
FIG. 20-54
(1 – 0.275)
4. Estimate vaporization losses from Fig. 20-55.
Weight % vs. Mol% for Methanol and EG Solutions
@ 4°C and 6200 kPa (ga), losses = (30 kg/106 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%
@ 4°C and paraffinic fluid, Dist. Ratio = 110
mol % MeOH in hyd. liquid = 17.5/110 = 0.16 mol%
70%
60%
1 m3 of condensate has a mass of 780 kg
50%
= (780/140) =5.6 kmol/m3 of condensate
40%
= (5.6)(0.0016) =0.009 kmol MeOH
30%
= (32)(0.009) = 0.29 kg/m3
20%
Total MeOH losses to the hydrocarbon liquid phase
= (0.29)(2.83)(56) = 46 kg MeOH/d
Methanol
10%
EG
Total methanol injection rate = 750 + 1490 + 46
= 2286 kg/d
0%
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
kg methanol per million Std m3
Mole percent methanol in aqueous liquid
10%
19.4oC
100
10.4oC
0.5 oC
19.4°C
15.3°C
5.0 oC
10.4°C
(
0%
Methanol in Vapor Phase
wt%
XL – XR
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
d = 14°C, M = 62, KH = 1297
Methanol Disadvantages
• Higher inhibitor losses to the hydrocarbon vapor and liquid phases
Solving for XI, XI = 0.40
2.Calculate mass rate of inhibitor solution in water phase
from Equation 20-9
• More difficult to recover methanol from the aqueous
phase
(0.40)(1980)
mI =
(0.80 − 0.40)
• More toxic than EG
• More flammable than EG (lower flash point)
= 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:
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
Methanol Advantages
• Generally less expensive than EG
• Much easier to recover from the water phase (regeneration)
• Requires lower concentrations in the aqueous phase
• Less toxic and less flammable than methanol
• Can inhibit to very low temperatures
• Can also provide corrosion inhibition for “top of the line
corrosion” in pipelines
• Has a lower viscosity than EG
FIG. 20-56
Liquid-Liquid Methanol Distribution Ratios
100
No Toluene
28-33mol% Toluene
Distribution Ratio
(
Mole fraction MeOH in Aqueous Phase
Mole fraction MeOH in Organic Phase
)
1000
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
• Increased production rates, where inhibitor injection capacity or flowline capacity is limited.
EG Disadvantages
• Higher concentrations required in the aqueous phase
• 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 the 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.
• 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
• 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.
• These are especially appropriate for offshore where
weight and space are critical to costs.
• Reduced operating expenses in many cases through decreased chemical consumption and delivery frequency.
20-29
FIG. 20-57
Hydrate Inhibition with Methanol and NaCI
11000
Pure Water
10000
20% NaCl
9000
20% MeOH
8000
Pressure, kPa
• Higher viscosity makes physical separation from hydrocarbon liquid phase more difficult
5% MeOH-15% NaCl
7000
15% MeOH-5% NaCl
6000
5000
4000
3000
2000
1000
0
0
4
8
12
Temperature, °C
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.
• The KHI polymer suffers degradation effects at temperatures above 249°C.
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.
• 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.
• 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.
­
—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.
• 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.
• 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.
• 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.
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 (see 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 TEG solution at various temperatures and TEG concentrations. Fig. 2059 can be used to estimate the required TEG concentration 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 H2O 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 gas 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 %
Dewpoint, °C
0.0
99.5 wt %
-10.0
99.7 wt %
-20.0
99.9 wt %
-30.0
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 wt%
1.00
Water Removal Fraction
0.90
N=3.0
N=2.5
0.80
N=2.0
0.70
N=1.5
0.60
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 pressures 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
N=3.0
0.90
N=2.5
0.85
N=2.0
0.80
0.75
N=1.5
0.70
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
Water Removal Fraction
Water Removal Fraction
0.95
N=4.0
N=3.0
0.90
N=2.5
0.85
N=2.0
0.80
0.75
N=1.5
0.70
0
10
20
30
40
50
Circulation Ratio, mass lean TEG sol/mass water absorbed
20-33
60
70
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.
The height of packing can be estimated from Equation 20-10.
Packing Height = (NTU)(HTU)
Eq. 20-10
FIG. 20-63
Absorber Performance at Lean TEG Concentration = 99.9 wt%
1.00
Water Removal Fraction
N=4.0
N=3.5
0.95
N=3.0
N=2.5
0.90
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
Water Removal Fraction
0.98
N=4.0
0.97
N=3.5
0.96
0.95
N=3.0
0.94
0.93
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 Fs 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.
For packing, from Fig. 20-65 at a circulation ratio = 28 kg
TEG/kg H2O the required NTUs = 4.1
Feed gas density =
Example 20-11 — 0.85 106 Sm3/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 H2O 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.
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
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:
G = C [ρv (ρL – ρv)]0.5
Assume a 6°C approach to equilibrium so the equilibrium
dew point is –10°C @ T = 38°C, lean TEG concentration ≈
98.9 wt% (use 99.0 wt%)
= 176 [32 (1120 – 32)]0.5 = 32 800 kg/m2 • h
1 kmol 
228 000 Sm   · =  m

  23.64 
d


 
3
2.Estimate number of theoretical stages.
Calculate water removal efficiency
From Fig. 20-4 the inlet gas water content at 38°C and
4100 kPa (abs) = 1450 kg/106 Sm3
=
1450–110
1450
= 0.924
 (0.65)(28.97)
kg  1 d 

  24 h  = 28 200 kg/h
kmol



·
m
28 200
A=
=
= 0.86 m2
G 32 800
FIG. 20-65
Transfer Units (NTUs) vs. Theoretical Stages (N)
12.0
10.0
8.0
NTUs
Win
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
= 32 kg/m3
From Fig. 20-66, HTU = 0.76 m.
Solutions Steps:
Win – Wout
(4100)(0.65)(28.97) (0.92)(8.314)(273+38)
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)
HTU, m
0.8
0.7
As = 300 m 2/m 3
=(91 ft 2/ft 3)
0.6
As = 350 m 2/m 3
=(107 ft 2/ft 3)
0.5
0.4
0
25
50
75
100
125
Gas Density, kg/m3
4A
D =  
π


0.5
(4)(0.86)
= 
3.14


0.5
FIG. 20-67
= 1.05 m
Contactor Sizing Parameters for Glycol Contactors
For Structured packing:
 176 
=  
 384 
(1.05) = 0.7 m
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
Packing
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.
3
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 H2S 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.
xi =
C factor, m/h
Bubble Cap Trays
 CBubble cap 0.5
D =   (DBubble cap)
C
 Structured packing 
0.5
K factor, m/s
* 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
P
1
exp A + B
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
A
B
C
D
E
H2S in EG/
Water
4.256
–1298
1.824
0.02604
0.003354
H2S in TEG/
Water
3.037
–1414
1.4058
0.02130
0.003354
CO2 in EG/
Water
6.125
–1232
1.287
0.01934
0.003354
CO2 in TEG/
Water
4.813
–1239
1.429
0.01617
0.003354
Sensible Heat:
Qs = m Cp ∆t
(1120 kg)  2.78 kJ 
=
 kg • °C 
m3


= 156 000 kJ/m3
Vaporization of Absorbed H2O:
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).
Qc = (0.25)(Qv) = 22 500 kg/m3
Total Duty Including 10% Heat Loss:
Qr = (156 000 + 90 000 + 25 500)(1.1)
= 295 000 kJ/m3
Total Duty Based on 0.85 Sm3/d of gas:
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
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.
 2260 kJ 
1 kg H2O  = 90 000 kJ/m3
=    0.025 m3 TEG
O
kg
H
2



Condenser Duty @ 25% Reflux Ratio:
A quick estimate of reboiler duty can be made using Equation 20-3.
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.
Qv = (ΔHvap)(ΔW)
Regenerator sizing requires establishing the reboiler duty
and, when high TEG concentrations are required, providing
sufficient stripping gas.
Q = (418 000) (Lg )
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 H2S in Triethylene Glycol at 3500 kPa vs.
Temperature, H2S Content of Gas Phase, and Water Content of TEG
10.000
H2S Solubility, std. m3H2S / m3 TEG solution
H2S
H2S
H2S
H2S
1.000
H2S
H2S
H2S
0.100
H2S
H2S
H2S
H2S
H2S
0.010
0.001
20
30
40
50
(200ºC – 150ºC)
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 m3/h of TEG solution/106 Sm3/d of
gas. BTEX solubility increases with increasing circulation ratio
FIG. 20-69
H2S Solubility, std. m3H2S / m3 TEG solution
Approximate Solubility of H2S in Triethylene Glycol at 7000 kPa vs.
Temperature, H2S Content of Gas Phase, and Water Content of TEG
H2 S
10.00
H2S
H2S
H2S
H2S
1.00
H2S
H2S
H2S
H2S
0.10
H2S
H2S
H2S
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, CO2 Content of Gas Phase, and Water Content of TEG
CO2 Solubility, std m 3 CO2 / m 3 TEG solution
CO2
CO2
CO2
CO2
1.00
CO2
CO2
CO2
CO2
CO2
CO2
0.10
0.01
20
30
40
50
60
Temperature, °C
20-38
70
80
90
100
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.
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 in Figs 20-72 & 20-73 (2100 and
6900 kPa (abs)) are the two pressures at which experimental
data was collected in GPA RR-131.
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, CO2 Content of Gas Phase, and Water Content of TEG
CO2
CO2
CO2
CO2
CO2
1.00
CO2
CO2
CO2
CO2
CO2
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)
0.4
0.8
1.2
1.6
2.0
2.4
50
2.8
60
60
Solid lines:
25°C (77°F)
Dashed lines: 35°C (95°F)
Dotted lines: 50°C (122°F)
50
40
40
30
30
20
20
10
10
0
0.00
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
BTEX Absorption, %
0.0
BTEX Absorption, %
CO2 Solubility, std m 3 CO 2 / m 3 TEG solution
10.00
FIG. 20-73
BTEX Absorption in TEG Contactors at 1000 psia (6895 kPa)
Circulation Ratio, (m3/h TEG)/(10 6 Sm 3 /d of gas)
0.0
0.4
0.8
1.2
1.6
2.0
2.4
Solid lines:
25°C (77°F)
Dashed lines: 35°C (95°F)
Dotted lines: 50°C (122°F)
45
35
30
30
25
20
20
BTEX Absorption, %
40
40
BTEX Absorption, %
2.8
50
50
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 m /kg H2O absorbed) •
(0.85 106 Sm3/d)(1450–110) kg H2O/106 Sm3
= 1.2 m3/h
24 h/d
3
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.
•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 11oC higher than the ambient dry
bulb air temperature which limits the effectiveness of
condensation in warm climates.
Circulation ratio = (1.2 m3/h)/(0.85 106 Sm3/d) = 1.4
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
•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.
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)
•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.
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.
•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.
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.
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.
Recycle
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 C5+
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/106Sm3 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% H2O. 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 CO2 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 CO2. 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:
els — alumina or silica gels manufactured and condiG
tioned to have an affinity for water.
lumina — a manufactured or natural occurring form of
A
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 H2S 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 (Al2O3). 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 fired 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 = B μV + Cρ V
L
2
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., Dminimum),
and select the nearest standard diameter (i.e. Dselected):
 4q 
Dminimum =  
60 V
 π max
·m
q =
60ρ
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
0.5
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)
7x14 Tyler mesh
3.2 mm/4.8 mm/6.4 mm
1.00
–68°C dew point
735
3–6 mesh or 5–8 mesh
0.92
5–10 ppmv
675–720
4–8 mesh or 8–12 mesh
0.96
0.1 ppmv (–101°C)
640–705
3.2 mm/1.6 mm pellets
1.00
0.1 ppmv
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
Beads
720
5x8 mesh
1.05
–51°C dew point
Beads
720
5x8 mesh
1.00
–51°C dew point
Shape
Bulk ­
Density,­
kg/m3
Alumina Alcoa F200
Beads
770
Activated Alumina UOP A-201
Beads
Mole Sieve Grace – Davison 4A
Beads
Molecular Sieve UOP 4A-DG
Extrudate
Mole Sieve Zeochem 4A
Silica Gel Sorbead® – R
Silica Gel Sorbead® – H
Silica Gel Sorbead®– WS
Desiccant
Particle Size
20-44
FIG. 20-78
Static Equilibrium Capacity vs. Relative Humidity
for Selected Solid Desiccants46
Obtain the corresponding superficial velocity, Vadjusted as follows:
2
 Dminimum

Vadjusted = Vmax  Eq 20-17
Dselected 

 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
UOPTM Adsorbents
4A-DG MOLSIV TM 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 the 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 =
Wr
(0.13)(CSS)(CT )
(Ss)(4)
LS =
π(D2)(bulk density)
Molecular sieve bulk density is 675–735 kg/m3 for spherical
particles and 640–705 kg/m3 for extruded cylinders.
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:
Eq 20-19
FIG. 20-80
Allowable Velocity for Mole Sieve Dehydrator
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.
Now the total bed pressure drop is checked. The ∆P/L for the
selected diameter, Dselected, is adjusted using Equation 20-14 or
the following approximation:
 Vadjusted 2
(ΔP/L)adjusted ≅ (7.5 kPa/m)  
Eq 20-21
Vmax


The result is multiplied times the total bed height (LS +
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 of water on bed)
kg


Eq 20-22
 1.0 kJ 
Qsi = (kg of sieve)   (Trg – Ti)
kg • K


Eq 20-23
 0.5 kJ 
Qst = (kg of steel)   (Trg – Ti)
kg • K


Eq 20-24
Qhl = (heat loss) = Qw + Qsi = Qst)(0.10)
Eq 20-25
Qw
transfers to the bed, vessel steel, and heat loss to atmosphere;
and the balance leaves with the hot gas.
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, Thot, is 28°C above the temperature, Trg, to which the bed must be heated. The temperature, Tb,
is the bed temperature at the beginning of regeneration, which is
the same as the dehydration-plant feed temperature.
The temperature, Trg, 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, Pdesign, 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 0.75Dbed
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
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.
•
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
calculated velocity is less than this, the regeneration gas rate,
· rg, must be increased by multiplying it by the ratio Vmin/V, and
m
the period of regeneration should be decreased by multiplying
it times the ratio V/Vmin. Equation 20-14 may also be used to
calculate Vmin using a ΔP/L equal to 0.23 kPa/m.
Eq 20-28
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
FIG. 20-83
Inlet and Outlet Temperatures During Typical Solid
Desiccant Bed Regeneration Period
20-47
FIG. 20-84
Minimum Regeneration Velocity for
Mole Sieve Dehydrator
­General Comments
Vadjusted = (758 m/h)(2.21/2.25)2 = 731 m/h (Equation
20-17)
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 Sm3). 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.
ẇ
= (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/m3. 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 Sm3/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
from Equation 20-14. Let the maximum ∆P/L be 7.5 kPa/
m.
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, Pdesign, be 110% of the operating
pressure:
Pdesign = (4140)(1.1) = 4554 kPa (ga)
z = 0.93 from Fig. 23-8 for 17.4 mole weight which is conservative
ρ = (18 mole weight) (4140 kPa (abs))/((8.3145)
(311.15 °K) (0.93)) = 31 kg/m3 (Equation 23-2)
(1000) (2.25) (4554)
t= = 40.3 mm (Equation 20-26)
260 000 – 1.2 (4554)
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)
μ = 0.015 mPa • s (Fig. 23-27)
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)
ṁ = [(2.85 • 10 Sm /day)/((24 h/day)(23.646
Sm3kmole))] (18 kg/kmole) = 90 400 kg/h
6
3
Qw = (4200 kJ/kg) (2004 kg water)
= 8 417 000 kJ (Equation 20-22)
q = (90 400 kg/h)/(31 kg/m )
= 2916 m3/h (Equation 20-16)
3
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)
20-48
6.Calculate the flow rate of regeneration gas using Equation 20-29. Let the heating time be 60% of the total regeneration period.
(60%) (12 hr) = 7.2 hours heating
ṁrg = (41 980 000 kJ)/((2.7 kJ/kg °C)
(288°C – 38°C) (7.2 h)) = 8638 kg/h
(Equation 20-29)
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.
Refluxing
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, 1μm filters are frequently installed at the outlet of the dehydration unit to protect
downstream equipment.
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.
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 CaCl2 hydrates (CaCl2 • XH2O). As
20-49
water absorption continues, CaCl2 is converted to successively
higher states of hydration — eventually forming a CaCl2 brine
solution.
10 to 20 mm CaCl2 pellets are installed in a fixed bed much
like a dry desiccant tower. Gas flow is upflow. The more efficient
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 solid CaCl2 and increases
unit capacity. One such unit is shown in Fig. 20-85.
The solid CaCl2 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 CaCl2 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 have been achieved
with CaCl2 dehydrators. Typical CaCl2 capacity is 0.3 kg CaCl2
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.
3
CaCl2 dehydrators may offer a viable alternative to glycol units on low rate, remote dry gas wells. The CaCl2 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, CO2 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
20-50
the feed gas and permeate stream, the surface area of the membrane and the membrane selectivity. For membrane systems
designed to remove CO2 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.
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.
•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 temperature 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 H2S, 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.
Refer to the discussion of molecular sieve for gas dehydration elsewhere in this Section for more information.
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.
­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.
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.
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).
­Calcium Chloride
10.Gillespie, P. C., and Wilson, G. M., RR-41, “Vapor-Liquid Equilibrium Data on Water-Substitute Gas components: N2-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+CO2, Ethane + CO2,” 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.
20-52
15.Song, K. Y., and Kobayashi, R., “Water Content of CO2 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 220oF 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.
24.Mehta, A. P., and Sloan, E. D. “Structure H Hydrates: The State
of-the-Art,” Proceedings 75th GPA Annual Convention, Denver,
CO, 1996.
25.Katz, D. L., “Prediction of Conditions for Hydrate Formation in
Natural Gases,” Trans. AIME Vol. 160, 1945, p. 140.
26.Carson, D. B., and Katz, D. L., “Natural Gas Hydrates,” Trans.
AIME Vol. 146, 1942, p. 150.
27.Robinson, D. B., and Ng, H. L., “Improve Hydrate Predictions,”
Hyd. Proc. Vol. 54, No. 12, Dec. 1975, p. 95.
28.Poettmann, F. H., “Here’s Butane Hydrates Equilibria,” Hyd.
Proc., Vol. 63, No. 6, June 1984, p. 111.
29.Sloan, E. D. et.al., “Vapor-Solid Equilibrium Ratios for Structure
I & II Natural Gas Hydrates,” Proceedings 60th GPA Annual
Convention, San Antonio, Tx., 1989.
30.McCleod, H. O., and Campbell, J. M., “Natural Gas Hydrates at
Pressures to 10,000 psia,” Trans. AIME, Vol. 222, 1961, p. 590.
31.Blanc, C., and Tournier-Lasserve, J., “Controlling Hydrates in
High Pressure Flowlines,” World Oil, Vol. 211, No. 5, p. 63.
32.Noaker, L. J., and Katz, D. L., “Gas Hydrates of Hydrogen Sulphide-Methane Mixtures,” Trans. AIME Vol. 201, 1954, p. 237.
46.Ng, H.-J., et al., RR-149, “Vapour-Liquid and Vapour-Liquid-Liquid Equilibria for H2S, 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.
56. R
osman, A., “Water Equilibrium in the Dehydration of Natural
Gas with Triethylene Glycol,” Trans. AIME, Vol. 255, 1973, p.
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64. R
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Association, May 1991.
57. H
erskowitz, M., and Gottlieb, M., “Vapor-Liquid Equilibrium in
Aqueous Solutions of Various Glycols and Poly (Ethylene) Glycols, Triethylene Glycol,” J. Chem. Eng. Data, Vol. 29, 1984, p.
173.
65. Moshfegian, M., and John M. Campbell & Co., Aug 2011.
66. L
arue, J., et al, IFP Internal Note JL/MZ No. 363, “Calculation of
methanol dissolution in TEG” based on Lichtkin et al., Tr Khim
Tekhnol, March 3, 1973.
58. W
on, K.W., “Thermodynamic Basis of the Glycol Dew-Point
Chart and its Application to Dehydration,” Proceedings 73rd
GPA Annual Convention, New Orleans, LA, 1994, pp 108–133.
67. W
allace, C. B., “Dehydration of Supercritical CO2,” Proceedings
1985 Gas Conditioning Conference, Univ. of Oklahoma, Norman,
OK.
59. B
estani, B, Shing, KS, “Infinite dilution activity coefficients of
water in TEG, PEG, glycerol and their mixtures in the temperature range 50 to 140ºC,” Fluid Phase Equilibria, vol 50, 1989.
68. K
ohl, A. L., and Riesenfeld, F. C., Reprinted from Figure 12–5,
page 637, “Gas Purification,” 4th edition, 1985, Gulf Publishing
Company, Houston, Texas, with permission from Elsevier.
60. H
ubbard, R. and Clinton, P.,” Estimation of Glycol Contactor
Performance Using Graphical Techniques,” Laurance Reid Gas
Conditioning Conference, Norman, Oklahoma, Feb 2009.
69. E
rgun, S. “Fluid Flow Through Packed Columns,” Chem. Eng.
Prog., Vol. 48, No. 2, Feb. 1952, p. 89.
70. M
inkkinen, A., Larue, J., and Patel, S., “Methanol Gas Treating
Scheme Offers Economics, Versatility,” The Oil and Gas Journal,
June 1, 1992, p. 65.
61. M
amroush D. Trimeric, “GPA PROJECT No. 975-8, CO2/H2S
Solubility in Glycols, An Update of Engineering Data Book Information,” A summary of data presented in RR-183 and RR-189,
2011.
62. D
avis, P. M., Clark, P. D., Fitzpatrick, E., Lesage, K. L., Svrcek,
W. Y., and Satyro, M., RR-183, “The Impact of Sulfur Species on
Glycol Dehydration — A Study of the Solubility of Certain Gases
and Gas Mixtures in Glycol Solutions at Elevated Pressures and
Temperatures,” Alberta Sulfur Research Laboratories, Calgary,
Alberta, Canada, Project 992-1, January 2003.
63. Davis,
P. M., Clark, P. D., Fitzpatrick, E., Lesage, K. L., Svrcek,
W. Y., and Satyro, M., RR-189, “The Impact of Sulfur Species on
Glycol Dehydration — A Study of the Solubility of Certain Gases
and Gas Mixtures in Glycol Solutions at Elevated Pressures and
Temperatures,” Alberta Sulfur Research Laboratories, Calgary,
Alberta, Canada, Project 992-2, December 2005.
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