spe 169444 method to size gas-liquid horizontal separators handling nonstable multiphase streams

8/9/2019 SPE 169444 Method to Size Gas-Liquid Horizontal Separators Handling Nonstable Multiphase Streams

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SPE-169444-MS

A Method To Size Gas-Liquid Horizontal Separators Handling NonstableMultiphase Streams

J.L. Hernandez-Martinez, Pemex; V. Martinez-Ortiz, Schlumberger

Copyright 2014, Society of Petroleum Engineers

This paper was prepared for presentation at the SPE Latin American and Caribbean Petroleum Engineering Conference held in Maracaibo, Venezuela, 2123 May

2014.

This paper was selected for presentation by an SPE program committee following review of information contained in an abstract submitted by the author(s). Contents

of the paper have not been reviewed by the Society of Petroleum Engineers and are subject to correction by the author(s). The material does not necessarily reflect

any position of the Society of Petroleum Engineers, its officers, or members. Electronic reproduction, distribution, or storage of any part of this paper without the written

consent of the Society of Petroleum Engineers is prohibited. Permission to reproduce in print is restricted to an abstract of not more than 300 words; illustrations may

not be copied. The abstract must contain conspicuous acknowledgment of SPE copyright.

Abstract

The sizing procedures of all the process equipment in oil and gas industry share a common assumption:

the stream flowrate measured at the inlet is always constant or stable. This principle works fine for plants

where the process conditions are controlled, but it does not apply to primary production facilities where

the streams come from multiphase flowlines or wells. In a real horizontal gas-liquid separator, the inlet

stream flowrate is not constant and the liquid fraction fluctuates as a function of time. This fact must be

considered in the design procedure applied to these units.

The sizing procedure described here allows to size horizontal gas-liquid separators, consideringnon-stable flowrate at the inlet, or fluctuations in the liquid fraction on the same stream. The method looks

for the best unit design to address these variations and also gets the smallest possible horizontal gas-liquid

separator.

The proposed method consists of two parts. The first one is solved by a dynamic multiphase flow

simulator which allows modeling the multiphase flow in the pipeline and predicting flowrate and liquid

fraction fluctuations at the separator inlet. In the second part, the horizontal separator sizing is solved

following a modifiedSvcek and Monnery (1993) method. It is adapted to take into account a non-stable

stream at the inlet, using the surge volume calculated by dynamic multiphase simulation. Other well-

known methods are also reviewed, such asArnold (1986). Relationships between the Svcek and Monnery

approach and other methods are depicted.

The method is illustrated sizing a gas-light oil separator located in an offshore production facility where

the feed stream is identified as non-stable. This behavior is taken into account, and the effect in the vessel

volume is numerically evaluated. This approach allows obtaining the smallest separator volume based on

dynamic simulation results; as compared to the sizing based on overdesign rules.

The approach depicted here can be applied for sizing horizontal gas-liquid separators in which the

dynamic nature of the multiphase flow must be taken into account and the smallest possible volume is

desired.


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Background

The sizing of separators is a common task on pro-duction facilities design. There are several proce-

dures published in the literature to size tanks [2,3,

5,and6]. All of them share the same basic princi-

ple: stream flowing to the separator is stable. This

assumption should be correct for equipment to be

installed in process plants, but for primary separa-

tors where the processed stream comes from wells

or multiphase flowlines, this assumption cannot be

applied. Multiphase streams exhibit a nonstable behavior due to terrain or hydrodynamic slugging. For

design purposes, the separator is split in three stages (or sections) [5]:

Primary stage. This is a mechanical operation; it uses an inlet diverter so that the momentum of

the liquid in the gas causes the largest droplets to impinge on the diverter, forming droplets which

are separated by gravity. The gas stream flows to the disengagement section.

Secondary stage. Liquid drops immersed in the gas stream are separated by gravity. This

separation is governed by the settling liquid drop velocity inside the gas phase. This velocity must

be higher than gas stream velocity to allow the drop settle to the liquid phase.

Tertiary Stage. It occurs on the mist eliminator. This is a device that allows small size droplets

coalesce to larger drops, which fall down by gravity to the liquid phase.

These design stages or sections are shown in Figure 1.

Primary stage design is based on a simple volumetric balance, following some well-known rules (to be

described below). The most important parameter to design the secondary stage is to obtain the settlingdrop velocity inside the gas stream. This velocity can be calculated performing a balance around each

liquid droplet. This balance takes into account two factors, the gravity and the drag forces applied by the

gas stream.Figure 2depicts both forces;FGandFDare gravity and drag components, respectively. Liquid

droplets settling at a constant velocity known as terminal velocity, calculated byEquation 3

Equation 1

If UV UTliquid droplets will settle. As a thumb rule, the allowable vertical velocity has a value

between 0.75 UTandUT. Equation 3can be rewritten as Sauders and Brown suggested[5].

Figure 2Forces over a liquid drop

Figure 1Scheme for a gas-liquid separator

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Equation 2

Where

Equation 3

Kis the terminal velocity constant. For practical reasons, very small droplets cannot be separated only

by gravity. Coalescent devices are used (mist eliminators) to force this separation. This constant,K, is not

easily predicted theoretically, empirical correlations are preferred. There are several published papers

depicting procedures to obtainK[1,3and6], but many of them should be applied to separators in process

plants not in primary oil-gas facilities.

Theoretical model to calculate terminal velocityTerminal velocity can be calculated for separators without mist eliminator using Equation 5, the drag

coefficient, CD

, is obtained usingEquation 6.

Equation 4

WhereDp is in feet, in lb/ft3 and in cP. Usually the drop diameter is given in microns (1 micron

3.28084 10-6 ft). A well-known sizing procedure was proposed by Arnold [2]. This approach is limited

to separators with a normal level equivalent to 50% of total vessel volume. An alternative procedure to

calculate the terminal velocity constant is proposed by this author. The terminal velocity constant iscalculated using curves shown inFigure 3.Constant gas viscosity and 100-micron droplets are assumed.

To avoid this graphical step,Equation 7is proposed to calculate this constant[4].

Equation 5

Where

Equation 6

Figure 3K values for separators (taken from Reference 2)

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Sis the relative gas density, P, the pressure in psia,APIis oil density andTthe temperature in Rankine.

K is calculated using Equation 7. Arnolds constant is a separator parameter, not a Sauders-Brown

constant. To convert Arnolds K to K, the following relationship is proposed [4].

Equation 7

Theoretical model (Equation 6) and Arnolds model have similar results.

Holdup and surge separator volumes

According to Svcek and Monnery [5], separator holdup is defined as the time it takes to reduce the liquid

level from normal (NLL) to empty (LLL) with a constant outlet drain rate, and a closed inlet. Holdup is

the capacity to keep a constant outlet flowrate interrupting the inlet flowrate. In other hand, surge is the

time it takes to raise the level from normal (NLL) to maximum (HLL), with a constant inlet flowrate and

a closed outlet stream. Surge is the capacity to accumulate liquid if the outlet stream is closed and the inletis open. There are several recommendations to state the holdup and surge times. These definitions work

fine for process vessels or balance tanks, but no for primary oil-gas separators. To apply Svcek and

Monnery procedure, some redefinitions are done. The holdup time is replaced by theresidence time. This

parameter assures gas-oil separation in the primary stage. Some suggestions to set the residence time are

shown inTable 1[1].

For foamy oils the residence time can be increased up to 15 min. Residence times between 2 and 5 min

are common for all the separators in the field [1].

Surge volume for nonstable multiphase streams

On primary separators where the feed comes from wells or multiphase flowlines, slugging can occur. As

a consequence, the feed flowrate exhibit a nonstable behavior, liquid fraction at separator inlet fluctuates

over the time. This fact does not agree the sizing procedure basic assumption, steady flow rate at the inlet.

To adapt Svcek and Monnery procedure to nonstable streams, surge volume is redefined. Surge volume

is the maximum liquid accumulation in the separator at constant outlet drain rate. This volume is reserved

to catch and damp the slugging occurring in the inlet stream. The objective is to design a separator

according to the residence time and disengagement gas section design rules. To perform the surge volume

calculations, a dynamic multiphase simulator is mandatory. A detailed description of liquid and gas

flowrates feeding the separator is required [4].

The maximum accumulation, defined as the surge volume, is given byEquation 8[4].

Table 1Residence time according to API 12 J

Residence (holdup) time

Oil density Time (min)

Above 34 API 1

Between 20 and 30 API 1.0 2.0

B etween 10 y 20 API 2.0 4.0

Figure 4Surge volume curve as a function of drain rate

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Equation 8

Where VSurge is the surge volume, , liquid accumulation measured at a fixed position (typically

flowline outlet), andQdrainis the constant outlet separator drain rate. Surge volume can be estimated based

on simulation results using an Excel spreadsheet. Optionally, this value can be calculated directly by some

codes. Surge volume usually is reported in curves as shown in Figure 4. Surge volumes make sense for

drain flowrates higher than the average liquid flowrate measured at flowline outlet.

Sizing procedure for nonstable multiphase streams

This procedure was initially described by Svcek and Monnery [5] and was applied to size separators in

process plants. The modification presented here adapts it to separators handling nonstable streams. Some

changes occur in the theoretical basis exposed previously. In horizontal separators, the liquid droplets are

influenced by a drag force in the horizontal direction; this force is not opposed to the gravity as in vertical

separators (seeFigure 2). As a consequence, the horizontal velocity is greater than the terminal velocity

(in the vertical direction). Liquid dropout time to go over the horizontal separator distance must be higher

than the settling time in the vertical direction.

Equation 9

Equation 9can be rewritten explicitly to the horizontal velocity.

Equation 10

Usually the L/HV is greater than one for horizontal separators and the terminal velocity must be

corrected. Theoretical or Arnolds models are recommended to calculate the terminal velocity constant.

Values obtained using these models are more conservatives than other cited in the literature [2,3 and6].

This sizing procedure is iterative; the separator diameter is not explicitly calculated. For this procedure the

design input data are the mass flowrates and fluid properties. These equations can be easily adapted to the

volumetric flowrates available in the field. Actual flowrates are required, no standard flowrates measured

at stock tank conditions. This rule must be applied to gas and liquid densities.

1. Calculate the volumetric gas flowrate.

Equation 11

2. Calculate the volumetric liquid flowrate.

Equation 12

3. Calculate the gas terminal velocity.UV 0.75 UT, for a conservative design (Equation 2).4. Calculate the holdup volume, based on residence time.

Table 2Typical L/D values

P (psig) L/D

0.0 250.0 1.53.0

250.0 500.0 3.04.0

500.0 4.06.0

Table 3Typical minimum level in horizontal separators

D (ft) LLL (in)

4.0 9.0

6.0 10.0

8.0 11.0

10.0 12.0

12.0 13.016.0 15.0

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Equation 13

5. Calculate the surge volume, based on dynamic simulation results (Equation 8).6. Obtain an estimate of separator diameter.

Equation 14

Table 2 shows typical values for L/D relationship to horizontal separators, as a function ofoperating pressure.Round to the nearest 6 inches. Calculate the cross-sectional area.

Equation 15

7. Calculate the minimum liquid level, based on Table 3.Alternatively,Equation 17can be used.

Equation 16

Where D is in feet and round to the nearest inch. IfD 4.0 ft, HLLL 9 in.8. Based onHLLL/D calculateALLL/D usingTable 4.9. CalculateHV. If there is no mist eliminator, the minimum height of the vapor disengagement area

(AV) is the larger of 0.2D and 1.0 ft. If there is mist eliminator the minimum height ofHVis thelarger 0.2 D and 2.0 ft. CalculateAVbased on HV.

10. Calculate the minimum length to contain holdup and surge volumes.

Equation 17

11. Calculate the dropout time.

Table 4 Cylindrical height and area conversions

Conversionh/Da/A

Y a/A, X h/D

a

4.755930

10

-5

b 3.924091

c 0.174875

d 6.358805

e 5.668973

f 4.018448

g 4.916411

h 1.801705

i 0.145348

Conversiona/A h/D

Y h/D, X a/A

a 0.00153756

b 26.787101

c 3.299201

d 22.923932

e 24.353518

f 14.844824

g 36.999376

h 10.529572

i 9.892851

Table 5Criteria to select vessel head

Criteria Heads

D15.0 ft, P100.0 psig Elliptical heads, 2:1

D15.0 ft for all P Hemispheric heads

D15.0 ft, P 100.0 psig Dished heads

Table 6Wall thickness

Head Thickness

Shell

Elliptical heads, 2:1

Hemispheric heads

Dished heads

Table 7Surface area

Head Surface area

Shell DL

Elliptical heads, 2:1 1.090D2

Hemispheric heads 1.571D2

Dished heads 0.842D2

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Equation 18

12. Calculate the actual gas velocity.

Equation 19

13. Calculate the minimum required length to gas-liquid disengagement section.

Equation 20

14. If LLMIN, thenL LMIN. Gas liquid separation is the controlling parameter, this result in someextra holdup volume.IfLMINL, the increase HVand repeat from step 9.IfLLMIN, the design is acceptable.IfLLMIN, holdup is controlling the design, L can only be decreased andLMINincreased ifHVis decreased. HV may only be decreased if it is greater than the minimum depicted in step 9.Calculations would be repeated from the step 9 with reducedHV.

CalculateL/D. IfL/D

6.0, increaseD and repeat from step 6. IfL/D

1.5, decreaseD and repeatfrom step 6.15. Calculate shell and heads thickness. Select heads type according to Table 5.

Calculate wall thickness according toTable 6.Pis the design pressure, in psig (typically 10 to 15 % above the operating pressure). D is thediameter in inches,Sthe material stress in psi, Eweld efficiency (typically 0.85) and y tCin theallowable corrosion, in inches. Wall thickness must be rounded to the nearest commercialthickness. Calculate head and shell surface according to Table 7.

16. Calculate the approximate vessel weight.

Figure 5Gas-liquid separator dimensions

Figure 6Gas-liquid separator dimensions (cross-area view)

Table 8Input data

Inl et te mperature (C) 79

Outlet pressure (kg/cm2) 59

Outlet pressure (C) 20

Oil f lowrate (BPD) 18,000

Gas flowrate (MMscfd) 17.7

Water flowrate (BPD) 0

GOR (Sm3/Sm3) 175

Density (API) 53.2

Length (km) 9.0

SPE-169444-MS 7


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Equation 21

Where ASandAHare shell and head areas respectively.17. Increment and decrease the diameter in 6 inches, repeat the calculation until L/D has ranged

between 1.5 and 6.0.18. With the optimum design (minimum weight) calculate normal and high liquid level.

Figure 7Pipeline scheme

Table 9Surge volumes for several drain rates

Case Average flowrate multiplier Surge volume (m3)

1 - 0.00

2 1.0 8.56

3 1.2 0.90

4 1.4 0.66

Figure 8Pressure at wellhead

Figure 9Total liquid flowrate at pipeline outlet

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Equation 22

Equation 23

Calculate ANLL based on HNLL.

Separator dimensions are shown in Figure 5 and Figure 6. This procedure was coded in an Excel

spreadsheet; it is available contacting the authors.

Case study

The procedure depicted here was used to size a separator for an oil field located in the Gulf of Mexico.

Table 8shows some general input data [4].

Figure 7shows a schematic diagram of this flowline. Separator is located at flowline outlet.

Results and discussion

A nonstable flowrate is observed in this flowline. Figure 8 shows wellhead pressure andFigure 9total

liquid flowrate measured at flowline outlet. Hydrodynamic slugging was identified by the dynamic

multiphase simulator. Trends observed inFigure 8andFigure 9confirms that assumption. A slug tracking

algorithm was used to calculate individual slugs generated in this flowline. Four cases were analyzed, in

the first one, the surge volume is neglected. In the additional cases, the surge volume was calculated fora constant drain rate.Table 9shows the surge volumes calculated for a drain rate 1, 1.2 and 1.4 times the

average total liquid flowrate, measured at separator inlet [4].

Separator sizes for the analyzed cases are shown inTable 10.

Volume increases around 52% if a drain rate equal to the average total liquid flow rate is used. This

increased volume is reduced if the drain rate is increased. The drain rate cannot be increased arbitrary; it

is function of the facilities downstream the separator. A thumb rule states that a drain rate 20% above the

average total liquid flow rate is acceptable for design purposes. If a drain rate 1.2 times the average total

liquid flowrate is used, total vessel volume increases 13%. If a 1.4 factor is used, the total vessel volume

increases 9%, this is a marginal reduction in the vessel size. A 1.2 factor seems to be fine for this design.

The separator sized following this procedure will handle efficiently the slugging generated in the

multiphase flowline.

Conclusions

A procedure to size gas-liquid separators, based on Svcek and Monnery approach, is depicted in this

paper. This modified procedure is adapted to size separators handling nonstable multiphase streams. This

procedure, replaces the original Svcek and Monnerys definitions of separator holdup and surge volumes,

by other characteristic parameters related to slugging. Surge volume is calculated using a dynamic

multiphase code, standard steady state tools cannot predict nonstable. Surge volume is a function of the

separator liquid drain rate.

Table 10Separator size for several surge volumes

Case

Length

(ft)

Diameter,

(ft)

Vessel volume

(ft3)

Vessel volume increase

(%)

1 20 4.5 318.1 -

2 28 5.5 665.2 52.2

3 23 4.5 365.8 13.0

4 22 4.5 349.9 9.1

SPE-169444-MS 9


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The procedure was applied to size a separator for a field located in Gulf of Mexico. If the slugging

effect is included in the sizing routine, an additional volume is required to catch and damp the slugs

generated in the flowline. In the studied case, if the drain rate is equal to the average total liquid flowrate,

the separator requires around of 52% extra volume. This additional volume decreases if the drain rate

increases. Vessel volume increases 13 and 9% if a drain rate 1.2 and 1.4 times the average total liquid

flowrate are used.

Nomenclature

CD Drag coefficient

D Diameter

DP Drop diameter

E Weld efficiency

F Force

g Gravity

G Gas-oil relationship

GOR

H Height

Hv Vertical separator distance

h Liquid level

j Time step

K Terminal velocity constant

K Separator constant

L Length

L Liquid phase

P Pressure

Qdrain Separator drain flowrate

S Gas relative density

S Gas relative density

S Maximum allowed stress

T Temperature

TH Residence time

T Time, thickness

tc Allowed corrosion

U Velocity

Ut Terminal Velocity

UV Allowable vertical velocity

V Volume, vapor section on sizing procedure

VSurge Surge volume

W Mass flowrate, total vessel weight Accumulative volume measured in a fixed position

Dropout time

G Gas viscosity

G Gas density

V Vapor density

References1 API, Specification for oil gas separators, API specification 12J, October 1989.

2 Arnold, K, M. Steward, Surface production operations, Gulf Publishing Co., 1986

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3 Gerunda, Arthur,How to size liquid-vapor separators, Chemical Engineering, May 1981.

4 Hernndez, Jos Luis,Dimensionamiento de separadores para lneas multifsicas con transporte

de flujo inestable empleando resultados de simulacin dinmica (in Spanish). Thesis, ESIA, IPN,

Mexico 2009.

5 Svcek, W, W, Monnery, Design two-phase separators within the right limits, Chemical Engi-

neering Progress, October 1993.

6 Watkins, R. N., Sizing separators and accumulators, Hydrocarbon Processing, November 1967

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spe 169444 method to size gas-liquid horizontal separators handling nonstable multiphase streams

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