deepstar_chevron-multiphase-design-guide-pt-1

PART I - MULTIPHASE PIPELINE & SLUG CATCHER DESIGN GUIDE

Multiphase Pipeline Design Guide

CPTC

NOVEMBER 1994

1


PART I

TABLE OF CONTENTS

SECTION 1.0 - INTRODUCTION 1.1

Objective and Scope..................................................................................................................................................... 1

1.2

Definition of Terms........................................................................................................................................................ 1

SECTION 2.0 – OVERVIEW OF MULTIPHASE FLOW FUNDAMENTALS 2.1

Design Criteria.............................................................................................................................................................. 11

2.2

Velocity Guidelines ....................................................................................................................................................... 11

2.3

Flow Regimes............................................................................................................................................................... 13

2.4

Pressure Gradient ......................................................................................................................................................... 16 2.4.1

Frictional Losses .......................................................................................................................................... 16

2.4.2

Elevational Losses........................................................................................................................................ 17

2.4.3

Acceleration Losses...................................................................................................................................... 18

2.4.4

Allowable Pressure Drop............................................................................................................................... 20

2.5

Pressure Gradient Calculations...................................................................................................................................... 20

2.6

Section Highlights......................................................................................................................................................... 21

SECTION 3.0 – STEADY STATE DESIGN METHODS 3.1

Pipeline Design Methods ............................................................................................................................................... 25

3.2

Steady State Simulators................................................................................................................................................ 26 3.2.1

Phase Equilibrium and Physical Properties.................................................................................................... 26

3.2.2

Pipeline Elevation Profile .............................................................................................................................. 28

3.2.3

Heat Transfer ............................................................................................................................................... 30

3.2.4

Recommended Methods for Pressure Drop, Liquid Holdup, and Flow Regime Prediction................................................................................................................................ 33

3.2.5 3.3

Interpretation of Results................................................................................................................................ 35


SECTION 4.0 – TRANSIENT FLOW MODELING 4.1

Transient Flow Modeling (General) ................................................................................................................................ 41

4.2

Use of Transient Simulators........................................................................................................................................... 42

CPTC

NOVEMBER 1994

2


4.3


SECTION 5.0 – SLUG FLOW ANALYSIS 5.1

Slug Flow (General) ...................................................................................................................................................... 45

5.2

Slug Length and Frequency Predictions......................................................................................................................... 46

5.3

5.2.1

Hydrodynamic Slugging................................................................................................................................ 46

5.2.2

Terrain Slugging........................................................................................................................................... 51

5.2.3

Pigging Slugs............................................................................................................................................... 53

5.2.4

Startup and Blowdown Slugs........................................................................................................................ 55

5.2.5

Rate Change Slugs ...................................................................................................................................... 56

5.2.6

Downstream Equipment Design for Slug Flow............................................................................................... 56


SECTION 6 – EXAMPLE PROBLEMS 6.1

6.2

Example Problem – 1 Low Gas/Oil Line Between Platforms .......................................................................................... 63 6.1.1

Line Size...................................................................................................................................................... 65

6.1.2

Slug Length Prediction ................................................................................................................................. 75

6.1.3

Slug Frequency and Length by Hill & Wood Method ...................................................................................... 80

Example Problem – 2 Gas Condensate Line .................................................................................................................. 88 6.2.1

Table 1, Wellstream Composition ................................................................................................................. 89

6.2.2

Table 2, Pipeline Evaluation Profile ............................................................................................................... 90

6.2.3

Pipeline Simulation Comparison ................................................................................................................... 92

SECTION 7.0 – REFERENCES .................................................................................................................................................... 106

FIGURES I: 1-1

Flow Regimes in Horizontal Flow................................................................................................................................... 8

I: 1-2

Flow Regimes in Vertical Flow ...................................................................................................................................... 9

I: 2-1

Horizontal Flow Regime Map......................................................................................................................................... 23

I: 2-2

Vertical Flow Regime Map............................................................................................................................................. 24

I: 5-1

Taitel-Dukler Liquid Holdup Predictions.......................................................................................................................... 60

I: 5-2

Stages in Terrain Slugging ............................................................................................................................................ 61

I: 5-3

Pipeline Slugging.......................................................................................................................................................... 62

I: 6-1

Liquid Holdup for Example 1, Year 12 ............................................................................................................................ 101

I: 6-2

Inlet Pressure for Example 1, Year 12............................................................................................................................ 102

I: 6-3

Liquid Flowrate Out of Line, Example 1, Year 12............................................................................................................ 103

I: 6-4

Gas Flowrate Out of Line, Example 1, Year 12............................................................................................................... 104

CPTC

NOVEMBER 1994

3


I: 6-5

CPTC

Liquid Holdup Predictions for Example 2........................................................................................................................ 105

NOVEMBER 1994

4


SECTION 1.0 - INTRODUCTION 1.1

Objective and Scope

The simultaneous flow of gas and liquid through pipes, often referred to as multiphase flow, occurs in almost every aspect of the oil industry. Multiphase flow is present in well tubing, gathering system pipelines, and processing equipment. The use of multiphase pipelines has become increasingly important in recent years due to the development of marginal fields and deep water prospects. In many cases, the feasibility of a design scenario hinges on cost and operation of the pipeline and its associated equipment. Multiphase flow in pipes has been studied for more than 50 years, with significant improvements in the state of the art during the past 15 years. The best available methods can predict the operation of the pipelines much more accurately than those available only a few years ago. The designer, however, has to know which methods to use in order to get the best results. Part I of this guide consists of seven sections. The fundamentals of multiphase flow in pipelines are discussed in Section 2.0. The third section describes the use of steady state simulation methods. This section of the guide helps the designer choose the best methods for the project, and it gives guidelines to use in designs. The fourth section of the report briefly describes transient flow modeling. The fifth section describes slug flow modeling, giving suggestions on the best methods to use in slug flow simulation. The sixth section includes two sample problems, based on actual designs, which illustrate the design steps used in analyzing the pipeline designs. 1.2

Definition of Terms

In discussing the design of multiphase pipelines, it is necessary to define several terms used repeatedly throughout this text. Near Horizontal and Near Vertical Angles The term "near horizontal" is used in this guide to denote angles of -10 degrees to +10

CPTC

NOVEMBER 1994

5


degrees from horizontal. The term "near vertical" denotes upward inclined pipes with angles from 75 to 90 degrees from horizontal. Flow Regimes In multiphase flow, the gas and liquid within the pipe are distributed in several fundamentally different flow patterns or flow regimes, depending primarily on the gas and liquid velocities and the angle of inclination. Observers have labeled these flow regimes with a variety of names. Over 100 different names for the various regimes and subregimes have been used in the literature. In this guide, only four flow regime names will be used: slug flow, stratified flow, annular flow, and dispersed bubble flow. Figure I:1-1 shows the flow regimes for near horizontal flow, and Figure I:1-2 shows the flow regimes for vertical upward flow. Descriptions of the flow regimes 1. Stratified Flow - at low flowrates in near horizontal pipes, the liquid and gas separate by gravity, causing the liquid to flow on the bottom of the pipe while the gas flows above it. At low gas velocities, the liquid surface is smooth. At higher gas velocities, the liquid surface becomes wavy. Some liquid may flow in the form of liquid droplets suspended in the gas phase. Stratified flow only exists for certain angles of inclination. It does not exist in pipes that have upward inclinations of greater than about 1 degree. Most downwardly inclined pipes are in stratified flow, and many large diameter horizontal pipes are in stratified flow. This flow regime is also referred to as stratified smooth, stratified wavy, and wavy flow by various investigators. 2. Annular Flow - at high rates in gas dominated systems, part of the liquid flows as a film around the circumference of the pipe. The gas and remainder of the liquid (in the form of entrained droplets) flow in the center of the pipe. The liquid film thickness is fairly constant for vertical flow, but it is usually asymmetric for horizontal flow due to gravity. As velocities increase, the fraction of liquid entrained increases and the liquid film thickness decreases. Annular flow exists for all angles of inclinations. Most gas dominated pipes in high pressure vertical flow are in annular flow. This flow regime is referred to as annular-mist or mist flow by many investigators.

CPTC

NOVEMBER 1994

6


3. Dispersed Bubble Flow - at high rates in liquid dominated systems, the flow is a frothy mixture of liquid and small entrained gas bubbles. For near vertical flow, dispersed bubble flow can also occur at more moderate liquid rates when the gas rate is very low. The flow is steady with few oscillations. It occurs at all angles of inclination. Dispersed bubble flow frequently occurs in oil wells. Various investigators have referred to this flow regime as froth or bubble flow. 4. Slug Flow - for near horizontal flow, at moderate gas and liquid velocities, waves on the surface of the liquid may grow to sufficient height to completely bridge the pipe. When this happens, alternating slugs of liquid and gas bubbles will flow through the pipeline. This flow regime can be thought of as an unsteady, alternating combination of dispersed bubble flow (liquid slug) and stratified flow (gas bubble). The slugs can cause vibration problems, increased corrosion, and downstream equipment problems due to its unsteady behavior. Slug flow also occurs in near vertical flow, but the mechanism for slug initiation is different. The flow consists of a string of slugs and bullet-shaped bubbles (called Taylor bubbles) flowing through the pipe alternately. The flow can be thought of as a combination of dispersed bubble flow (slug) and annular flow (Taylor bubble). The slugs in vertical flow are generally much smaller than those in near horizontal flow. Slug flow is the most prevalent flow regime in low pressure, small diameter systems. In field scale pipelines, slug flow usually occurs in upwardly inclined sections of the line. It occurs for all angles of inclination. Investigators have used many terms to describe parts of this flow regime. Among them are: intermittent flow; plug flow; pseudo-slug flow, and churn flow. Superficial Velocities The velocities of the gas and liquid in the pipe are prime variables in the prediction of the behavior of the multiphase mixture. Most multiphase flow prediction methods use the superficial gas and liquid velocities as correlating parameters. The superficial velocities are defined as the in situ volumetric flowrate of that phase divided by the total pipe crosssectional area. In oil field units, this corresponds to:

CPTC

NOVEMBER 1994

7


Vsg

= Superficial Gas Velocity, ft/sec = (actual ft3/sec of gas) / (pipe cross-sectional area, ft2)

Vsl

= Superficial Liquid Velocity, ft/sec = (actual ft3/sec of liquid) / (pipe cross-sectional area, ft2)

Mixture Velocity The mixture velocity (Vm) is the volumetric average velocity of the gas-liquid mixture. It is equal to the sum of the gas and liquid superficial velocities. Vm = Vsg + V sl

Slip and Liquid Holdup Liquid holdup is defined as the volume fraction of the pipe that is filled with liquid. It is the most important parameter in estimating the pressure drop in inclined or vertical flow. It is also of prime importance in sizing downstream equipment, which must be able to operate properly when the liquid holdup in the line changes because of pigging or rate changes. If there was no slip between the gas and liquid phases, both phases would move through the pipe at the mixture velocity. The liquid would occupy the volume fraction equivalent to the ratio of the liquid volumetric flowrate to the total volumetric flowrate. In multiphase flow terminology, this equates to the liquid holdup being equal to the ratio between the superficial liquid velocity and the mixture velocity: Hlns = No-slip liquid holdup = Vsl / Vm Under most conditions, however, the liquid phase, which is more dense and viscous, moves more slowly than the gas. When this occurs, the liquid holdup (Hl) is greater than the no-slip holdup.

CPTC

NOVEMBER 1994

8


H l > H ln s Under these conditions, the actual gas velocity is greater than the mixture velocity, and the actual liquid velocity is smaller than the mixture velocity. The expressions for the actual gas velocity (Ug) and actual liquid velocity (Ul) are: Ug =

Vsg 1 − Hl

V U l = sl Hl For small diameter, low pressure piping, there is frequently a vast difference between Ug and Ul. For field piping, there is generally less slip between the phases, and the flow may approximate no-slip flow in dispersed bubble and annular flows. It is possible to get conditions where the liquid holdup is less than no-slip, but this only occurs over a small range of flowrates in downwardly inclined pipes. Pressure Gradient Two definitions of the term "pressure gradient" are used in the literature. In this guide, the term "pressure gradient" will be used to describe the pressure drop per unit length of pipe, (Pin - Pout)/L. In many papers, the term "pressure gradient" is used to denote the pressure change per unit length (dp/dx = (Pout - Pin)/L). The magnitude of the pressure gradient is the same in either definition, but the sign of the pressure drop per unit length is usually positive, while the sign of dp/dx is usually negative. Most people prefer to work with positive numbers, so the majority of people use the pressure drop per unit length definition. The choice of the definition is somewhat arbitrary, but it should be noted when reading the multiphase flow literature, and working with some of the available software. 3-Phase Flow vs. 2-Phase Flow In most of this guide, the discussion will consider 2-phase flow, or gas-liquid flow. In the majority of oil field applications, there will actually be 3 phases present (gas, oil, and water). The rigorous prediction of 3-phase flow is in its infancy. 3-Phase flow methods

CPTC

NOVEMBER 1994

9


are not generally available, so most simulators use 2-phase models with a mixed liquid stream using averaged properties for the oil and water. The use of 2-phase models with averaged properties generally gives acceptable results unless either: emulsions are present; or the flowrates are low enough to cause stratification of all three phases. These problems are discussed in more depth in Section 3.2.1 Mechanistic Models vs. Correlations The prediction of multiphase flow behavior has improved considerably during the 50+ years that the subject has been studied. For many years, multiphase flow prediction methods were correlations, based on curve fits of experimental data. The correlations frequently used arbitrarily selected variables and were based on limited databases, consisting almost entirely of low pressure, small diameter data. Extrapolations of these prediction methods to field conditions frequently proved to be seriously in error. In the 1960s and 1970s, several investigators undertook experimental studies to try to understand the fundamental mechanisms of the various flow regimes. In the past 15 years models have been developed, which are based on simulation of these mechanisms. These models, referred to as mechanistic models, have proven to extrapolate best to field conditions. Newtonian vs. Non-Newtonian Fluids Most condensates and crude oils follow Newton’s law of viscosity, which is defined as:

τ yx = µ where τyx

dv x dy

= shear stress

µ

= viscosity

vx

= velocity

y

= distance

Some liquids, however, exhibit behavior that is very different from Newton's law. These fluids are referred to as non-Newtonian. In the oil field, examples of non-Newtonian fluids are drilling muds, polymeric additives, and crude oils below their cloud point.

CPTC

NOVEMBER 1994

10


Flowline simulators are based on Newtonian fluids. If a non-Newtonian liquid is present, the simulator must be “tricked” into giving a Newtonian viscosity equivalent to the actual viscosity at the given temperature and shear stress. The methods of doing this are beyond the scope of this guide.

CPTC

NOVEMBER 1994

11


Figure I:1-1 Flow Regimes in Horizontal Flow

CPTC

NOVEMBER 1994

12


Figure I:1-2 Flow Regimes in Vertical Flow

CPTC

NOVEMBER 1994

13


SECTION 2.0 - OVERVIEW OF MULTIPHASE FLOW FUNDAMENTALS 2.1

Design Criteria

The majority of lines are sized by use of three primary design criteria: available pressure drop; allowable velocities; and flow regime. In some cases, a more optimal line size may be found that better suits the overall design of the pipeline system. These considerations will be discussed later in the transient modeling section of the guide. A description of each of the primary design criteria follows in Sections 2.2, 2.3, and 2.4. 2.2

Velocity Guidelines

The velocity in multiphase flow pipelines should be kept within certain limits in order to ensure proper operation. Operating problems can occur if the velocity is either too high or too low, as described in the following sections. It is difficult to accurately define the point at which velocities are "too high" or "too low". This section of the guide will try to quantify limits, but these limits should be considered as guidelines and not absolute values. Maximum Velocity For the maximum design velocity in a pipeline, API RP-14E recommends the following formula: Vmax =

C ρ ns

(Eqn. 2.1)

where Vmax = Maximum mixture velocity, ft/sec

ρns

= No-slip mixture density, lb/ft3 =

ρg

CPTC

(ρ

g

)

Vsg + (ρ l Vsl ) Vm

= Gas Density, lb/ft3

NOVEMBER 1994

14


ρl

= Liquid Density, lb/ft3

C

= Constant, 100 for continuous service, 125 for intermittent service.

This equation attempts to indicate the velocity at which erosion-corrosion begins to increase rapidly. Many people think this equation is an oversimplification of a highly complex subject, and as a result, there has been considerable controversy over its use. For wells with no sand present, values of C have been reported to be as high as 300 without significant erosion/corrosion. For flowlines with significant amounts of sand present, there has been considerable erosion-corrosion for lines operating below C = 100. The use of the API equation has been the subject of several research projects. It has been generally agreed that the form of the equation is not sophisticated enough, and should include additional parameters. Unfortunately, no other equation has been proposed which has gained acceptance in the industry as an alternative to the API equation. As a result, the recommended maximum velocity in the pipeline is the value calculated from Equation 2.1 with a C value of 100. It should be noted that Equation 2.1 is also used by many people as an estimate of the maximum velocity for noise control. For additional guidance on the use of the API equation, refer to Chevron’s Piping Manual. Minimum Velocity The concept of a minimum velocity for the pipeline is an important one and should be considered in the design of the line. Turndown conditions frequently govern the design of the downstream equipment. Velocities that are too low are frequently a greater problem than excessive velocities, so that the designer’s natural tendency to add "a bit of fat" to the design by increasing pipe diameter can cause severe problems in the operation of the line and the downstream facilities. At low velocities, several operating problems may occur: a) Water may accumulate at low spots in the line. If there is an appreciable amount of CO2 or H2S in the well stream, this water may be very corrosive.

CPTC

NOVEMBER 1994

15


b) Liquid holdup may increase rapidly at low mixture velocities. A large accumulation of liquid may cause problems in downstream separators or slug catchers if the line is pigged or the rate is changed rapidly. c) Low velocities may cause terrain induced slugging in hilly terrain pipelines and pipeline-riser systems. It isn’t possible to give a simple formula quantifying the velocity when the phenomena discussed above will occur. The minimum velocity depends on many variables, including: topography; pipeline diameter; gas-liquid ratio; and operating conditions of the line. A ball-park value for the minimum velocity would be a mixture velocity of 5-8 ft/sec. The actual value of the minimum velocity can only be quantified by simulation of the system using the methods discussed in Section 5.2.2. 2.3

Flow Regimes

As discussed in Section 1, the gas and liquid in the pipe are distributed differently in each of the four major flow regimes (stratified, annular, slug, and dispersed bubble flows). The prediction of the correct flow regime is important for several reasons. The flow regime prediction can show whether the line will operate in a stable flow regime or an unstable regime. The prediction of liquid holdup and pressure drop are highly dependent on the flow regime, with each regime exhibiting different behavior when the design variables are changed. The transitions between the flow regimes are frequently depicted in a flow regime map, such as that shown in Figure I:2-1. The flow regime map typically has the superficial gas velocity (Vsg) on the X-axis and the superficial liquid velocity (Vsl) on the Y-axis. As discussed later in this section, the flow regime map is only valid for a single point in the pipeline. As the angle of inclination, pressure and temperature change with position in the pipeline, the flow regime map also changes. Some general comments, however, can be made about the flow regime transitions. Stratified flow occurs at low superficial gas and liquid velocities. Dispersed bubble flow occurs at high superficial liquid velocities. Annular flow occurs at high superficial gas

CPTC

NOVEMBER 1994

16


velocities. Slug flow occurs at moderate superficial gas and liquid velocities. Figure I:2-2 shows a typical flow regime map for vertical flow. Many experimental studies of the transitions between the flow regimes for various systems have been made, and many flow regime transition prediction methods have been published. Some of these methods work fairly well, but most are poor. The designer needs to carefully choose the method that will work best for the set of conditions. The best methods are discussed in the remainder of this section. Experimental studies of flow regime transitions have shown that each of the flow regime boundaries reacts differently to changes in the system variables. The following table shows the sensitivity of the transitions to changes in the major system variables:

Slug to Dispersed Bubble

Slug to Annular

Slug to Stratified

Stratified to Annular

Angle of Inclination

Small Effect

Moderate Effect

Strong Effect

Strong Effect

Gas Density

Small Effect

Strong Effect

Strong Effect

Strong Effect

Pipeline Diameter

Small Effect

Small Effect

Strong Effect

Moderate Effect

Liquid Physical Properties

Moderate Effect

Small Effect

Moderate Effect

Moderate Effect

Transition Variable

Many people have attempted to develop simple flow regime maps, usually using some arbitrary dimensionless parameter on each axis (e.g. Baker, Beggs & Brill). These methods are inherently inaccurate since no single parameter can model the sensitivity effects shown in the previous table. The only flow regime map prediction methods that have been effective for a wide range of conditions are those using mechanistic models to estimate the flow regime transitions.

CPTC

NOVEMBER 1994

17


In 1976, Taitel and Dukler published a landmark article describing a method of predicting flow regime transitions by modeling the mechanism of each transition. By modeling each transition, this method can show the same type of behavior observed in the experimental work. The original Taitel-Dukler paper covered flow regime transitions in near horizontal flow only, and one of the transitions (slug-dispersed bubble) is very much in error. Taitel and his co-workers at the University of Tel Aviv have subsequently published several articles that expand the range of angles of inclination and correct the errors in the original paper. The Taitel-Dukler paper and the latest paper from Tel Aviv model flow regime transitions for all angles of inclination. The Taitel, et al. methods give reasonably good predictions of the various flow regime transitions, and the accuracy of the predictions has improved with each revision. Another approach to the modeling of flow regime transitions is the method used in the OLGAS method. It employs mechanistic models of each flow regime and links the models by the assumption that the flow regime giving the lowest liquid holdup is the correct one. This assumption holds up well in practice. The OLGAS method predicts flow regime transitions with similar accuracy to the Taitel, et al. models. Within Chevron, there are several programs available for flow pattern prediction. Pipephase will print a flow regime map based on the Taitel-Dukler method for near horizontal flow and the Taitel-Dukler-Barnea model for near vertical flow. Unfortunately, these methods are the oldest and weakest of this family of methods. Two programs are available within CPTC that incorporate the latest versions of the Taitel, et al. models. These programs are FLOPAT, developed by Tulsa University, and FLEX, developed by Advance Multiphase Technology. CPTC should be consulted if it is desired to use these programs. As in many aspects of multiphase flow, the flow regime prediction methods are not exact. Errors of +/- 25% for the transition velocities are typical, even for the best prediction methods. If the Taitel-Dukler map is used, the designer should be aware of the gross errors in the slug to dispersed bubble transition. The errors for this transition can be 1000%. The dispersed bubble to slug transition typically occurs at a superficial liquid velocity of about 10 ft/sec. Taitel-Dukler frequently predicts this transition velocity to be 50-100 ft/sec.

CPTC

NOVEMBER 1994

18


2.4

Pressure Gradient

In most pipelines, the pipeline diameter is determined by the allowable pressure drop in the line. The overall pressure gradient is composed of three additive elements: a) pressure drop due to friction;

b) pressure changes due to elevational effects; c) accelerational losses. The calculation of the constituent parts of the pressure gradient will be discussed in the next three sections. The Chevron Fluid Flow Manual contains a good discussion of these pressure loss terms for single phase flow and can be consulted as a reference. 2.4.1

Frictional Losses

In multiphase flow, frictional losses occur by two mechanisms: friction between the gas or liquid and the pipe wall; and frictional losses at the interface between the gas and liquid. The friction calculations, therefore, are highly dependent on the flow regime, since the distribution of liquid and gas in the pipe changes markedly for each regime. In stratified flow, there is wall friction between the gas and the pipe wall at the top of the pipe, and wall friction between the liquid and the wall at the bottom of the pipe. There is also friction between the gas and liquid at the gas-liquid interface. The interfacial friction can be similar in magnitude to the wall friction if the interface is smooth, or it can be considerably higher if waves are present. In annular flow, there is friction between the liquid film and the wall. There is also considerable interfacial friction between the gas in the core of the pipe and the liquid film. The interfacial friction is usually the larger component. In dispersed bubble flow, friction occurs between the liquid and the wall. There is negligible interfacial friction between the gas and liquid.

CPTC

NOVEMBER 1994

19


Slug flow has several frictional components. In the slug, the friction losses are caused by the friction between the liquid and the pipe wall. In the gas bubble, the frictional components are the same as in stratified flow, namely gas and liquid friction with the pipe walls and interfacial friction between the gas and liquid. The friction loss in the slug is usually much higher than the losses in the bubble. 2.4.2

Elevational Losses

Elevational losses may be the major pressure loss component in vertical flow and flow through hilly terrain. The calculation of elevational losses is governed by the following equation:

ρ g sinα  dp    = mix  dx  elev 144g c where: (dp/dx)elev = Pressure gradient due to elevation, psi/ft

ρmix

= Mixture Density, lb/ft3 = (ρl) (Hl) + (ρg) (1-Hl)

Hl

= Liquid Holdup

g

= Acceleration due to gravity, 32.2 ft/sec2 α

gc

= Angle of inclination = Gravitational conversion factor, 32.2 lb-ft/(lbf-sec2)

In order to calculate the elevational gradient, the liquid holdup must be determined. The holdup in each flow regime has its own sensitivity to the important operating variables. A summary of the effect of the major operating variables on the liquid holdup is:

CPTC

NOVEMBER 1994

20


Slug Flow

Annular Flow

Stratified Flow

Dispersed Bubble Flow

Superficial Gas Velocity

Strong

Strong

Strong

Strong

Superficial Liquid Velocity

Strong

Strong

Strong

Strong

Gas Density

Moderate

Strong

Strong

None

Pipeline Diameter

Moderate

Weak

Weak

Weak

Angle of Inclination

Moderate

Weak

Very Strong

None

Liquid Properties

Moderate

Moderate

Moderate

Weak

As seen in the previous table, the influence of the major variables on the holdup is very different for each of the flow regimes. As a result, it is impossible to develop a general holdup correlation that will apply to all the flow regimes. Unfortunately, almost all of the commonly used holdup methods available in commercial software try to do this. They work poorly over much of the operating range as a result. The only way to accurately predict liquid holdup is to use mechanistic models for each of the flow regimes. The accuracy of available holdup methods is discussed further in Section 3.2.4. 2.4.3

Acceleration Losses

Although acceleration losses are present for all flow regimes, they are only significant for two flow regimes: annular flow and slug flow. The mechanisms for the losses in these two flow regimes are very different and will be discussed separately. In single phase flow, acceleration losses can be calculated from Bernoulli’s equation. Acceleration losses represent the change in kinetic energy as the fluid flows down the pipe. The expression for acceleration gradient is:  ρ V   dV   dp     =   dx  accel  144 g c   dx 

CPTC

NOVEMBER 1994

21


where: ρ

= Density, lbm/ft3

V

= Velocity, ft/sec

For multiphase flow, the same type of relationship holds except that it refers to the flow of the mixed phase fluid. Most methods assume a no-slip mixture and use the no-slip mixture density (ρns) and the mixture velocity (Vm) in the calculation of acceleration losses. The kinetic energy acceleration losses are small for most oil industry applications. The main exception is high velocity flow through low pressure piping. Flare systems would be an example of piping that has high acceleration losses. Acceleration may account for 3050% of the overall pressure loss in such lines. For a typical high pressure gathering system line, acceleration is usually less than 1% of the total drop and is frequently ignored. Please note that the present version of Pipephase, 6.02, does not properly account for acceleration losses, and, as a result, is not suitable for use in flare system design. In slug flow, another source of acceleration contributes significantly to the total pressure drop. As a slug propagates down the pipeline, it overruns and entrains the slower moving liquid from the film ahead of the slug front. Accelerating the liquid from the film velocity to the slug velocity can produce significant pressure losses. The acceleration loss may be anywhere from <1% to more than 50% of the total pressure drop. Mechanistic models include this loss, while most of the correlation based methods ignore this loss. 2.4.4

Allowable Pressure Drop

No clear-cut criteria exist for determining the amount of pressure drop to be allowed in a pipeline design. Allowable pressure drop is a function of the parameters of the system being designed. The following are some guidelines for specific systems: a)

For plant piping, rule of thumb values for pressure gradients, such as a frictional gradient of 0.2-0.5 psi per 100 ft. of equivalent length, are generally used.

b)

In the design of a gathering system, the ideal way to choose allowable pressure drops is to simulate the system from the reservoir through the processing plant as a function of time. This approach will account for the changes in reservoir pressure,

CPTC

NOVEMBER 1994

22


flowrate, and composition that the gathering system must handle during the life of the field. c)

If it isn’t feasible to do the rigorous simulations for a gathering system, the allowable pressure drop can be estimated from the initial wellhead pressure and the processing plant inlet separator pressure. A rule of thumb to use for this method is to take 1/3 of the difference between the wellhead pressure and the separator pressure as the allowable pressure drop in the pipeline. The remainder of the difference would equal the initial choke pressure drop. This approach would allow for future operation at reduced reservoir pressures.

d)

A rule of thumb estimate of allowable pressure drop for long distance gas/condensate pipelines is to allow 10-20 psi per mile for frictional pressure drop at design rates.

2.5

Pressure Gradient Calculations

As indicated in sections 2.4.1 to 2.4.3, the calculation of the pressure gradient for multiphase flow is very complicated. Hundreds of methods have been proposed to predict pressure drops, but only a few methods work well over a wide range of conditions. The best available methods are discussed in Section 3.2.4.

CPTC

NOVEMBER 1994

23


2.6

Section Highlights

Points to remember from Section 2.0 •

No other equation has gained acceptance in the industry like the API equaton. The recommended maximum velocity in the pipeline is the value calculated from Equation 2.1 with a C value of 100.

•

The Taitel et al. Methods give reasonably good predictions of the various flow regime transitions. The accuracy of the predictions has improved with each revision. The OLGAS method predicts flow regime transitions with similar accuracy to the Taitel et al. methods. If the Taitel-Dukler map is used, the designer should be aware of the gross errors in the slug to dispersed bubble transiton. Overall pressure gradient is composed of three additive elements: − pressure drop due to friction − pressure changes due to elevational effects − accelerational losses Frictional calculations are highly dependent on the flow regime, since the distribution of liquid and gas in the pipe changes markedly for each regime. Elevational losses may be the major pressure loss component in vertical flow and flow through hilly terrain. Using mechanistic models for each flow regime is the only way to accurately predict liquid holdup. Kinetic energy acceleration losses are small for most oil industry applications. The main exception is high velocity flow through low pressure piping. Pipephase 6.02 does not properly account for acceleration losses and is not suitable for use in flare system design as a result. For plant piping, rule of thumb values for pressure gradients, such as a frictional gradient of 0.2-0.5 psi per 100 ft. of equivalent length, are generally used.

• • •

• • • • • • •

CPTC

The allowable pressure drop for a gathering system can be estimated from the initial wellhead pressure and the processing plant inlet separator pressure. The rule of thumb for this method is to take 1/3 of the difference between the wellhead pressure and the separator pressure as the allowable pressure drop in the pipeline.

NOVEMBER 1994

24


•

CPTC

The rule of thumb for estimating allowable pressure drop for long distance gas/condensate pipelines is to allow 10-20 psi per mile for frictional pressure drop at design rates.

NOVEMBER 1994

25


Figure I:2-1 Horizontal Flow Regime Map

CPTC

NOVEMBER 1994

26


Figure I:2-2 Vertical Flow Regime Map

CPTC

NOVEMBER 1994

27


SECTION 3.0 - STEADY STATE DESIGN METHODS

3.1

Pipeline Design Methods

As stated in the previous sections, the pipeline designer needs to estimate the pressure drop, flow regime, and velocities in the line in order to select the proper line size. The calculation of these parameters is laborious and is usually done by computer simulation. Line sizing is usually performed by use of steady state simulators, which assume that the pressures, flowrates, temperatures, and liquid holdup in the pipeline are constant with time. This assumption is rarely true in practice, but line sizes calculated from the steady state models are usually adequate. Within Chevron, Pipephase and PIPEFLOW-2 are available for steady state pipeline simulation. For a more rigorous pipeline sizing, the simulations could be done using transient simulators. Transient simulators allow changes in parameters such as inlet flowrate and outlet pressure as a function of time, and calculate values for the outlet flowrates, temperatures, liquid holdup, etc. as a function of time. If the line is operating in slug flow, the line size calculated from the transient model may be different from that calculated from a steady state simulator. The principal uses of transient simulators are in the design of downstream equipment and the development of operating guidelines. Transient simulators can model transient behavior such as slug flow, pigging, rate changes, etc. Transient simulators are quite new, developed in the last 10 years, and are not in general use. Chevron has used the OLGA program for transient flowline analysis on several projects, utilizing outside consulting services. CPTC developed an in-house transient simulator, but it currently does not have as many features as the commercially available codes. The use of steady state models will be further discussed in Section 3.2, and transient modeling will be briefly discussed in Section 4.

CPTC

NOVEMBER 1994

28


3.2

Steady State Simulators

This section contains some general guidelines on the use of steady state simulators. Although there are several steady state programs available, the discussion will center on the use of Pipephase, which is Chevron’s currently recommended simulator. The topics covered include: a)

Phase Equilibrium and Physical Properties

b)

Pipeline Elevation Profile

c)

Heat Transfer

d)

Recommended Methods for Pressure Drop, Liquid Holdup, and Flow Regime Prediction

e)

Interpretation of Results

3.2.1

Phase Equilibrium and Physical Properties

Accurate prediction of the phase behavior and physical properties for the fluid flowing through the pipeline is essential to a good simulation of the pipeline operation. The estimates of these parameters depend in large part on the quality of the input data available. During conceptual design work, the only data available may be an estimate of the oil rate and gas-oil ratio. After well tests have been performed, compositions of the wellstream and PVT data may be available as well as projections of the flowrates of oil, gas and water as a function of time. Obviously, as the accuracy of the input data improves, the quality of the pipeline simulation improves. Pipephase has two fundamentally different models available within it for estimation of phase behavior and physical properties. The black oil model estimates the phase behavior and physical properties by use of a series of correlations that are based on operating temperature, pressure and some global parameters such as specific gravity of the oil and gas. Compositional models use an equation of state to estimate the quantity of liquid and gas at the operating conditions; then, correlations are used to estimate the physical properties. The decision on whether to use the black oil model or compositional modeling depends on the available information and the type of system that is being modeled.

CPTC

NOVEMBER 1994

29


The choice of models for gas-condensate and volatile oil systems is clear. Compositional models should be used for any gas-condensate or volatile oil system. This recommendation covers gas-oil ratios above about 3500 SCF/bbl. For lower gas-oil ratios, the choice of models is more difficult. Compositional models should give more accurate phase equilibrium results, but the physical property estimates from the compositional models may not be as good as the black oil model. (Section 6-1 illustrates this point.) As a result, it cannot be stated categorically that either the black oil model or the compositional model is superior for low gas-oil ratio systems. General practice with Pipephase has been to use the black oil model for lower gas-oil ratio streams. The accuracy of compositional modeling depends, in a large part, on the characterization of the heavy ends of the well stream. The materials heavier than hexane (C6+) are usually characterized by use of pseudo-components or cuts. The heavy ends could be characterized by one C6+ cut, or by a series of cuts corresponding to various boiling ranges. In general, the accuracy of the predictions increases when more cuts are used. Pipephase requires two of the following parameters in order to characterize a cut: specific gravity; molecular weight; or normal boiling point. In many cases, the mole fractions for cuts heavier than C6 may have been measured in the PVT analysis, but cut properties were not noted. In cases like this, the customary assumption is to use the properties of the corresponding normal paraffin as the cut properties. This adds some error to the analysis, but it is unavoidable in many circumstances. If tests of the phase equilibrium and physical properties have been done as part of the wellstream analysis, Pipephase allows the users of the black oil model to adjust the model predictions for solution GOR, densities, and liquid viscosity to match experimental values. The pipeline predictions after PVT matching should be considerably better than those obtained with use of the standard correlations. If the compositional model is used in Pipephase, the only variable that can be easily manipulated to match experimental data is the liquid viscosity. Pipephase does not have an option that will automatically adjust the phase equilibrium calculations to match experimental data. It is possible to manually modify the phase equilibrium calculations, but

CPTC

NOVEMBER 1994

30


it requires considerable effort, and the methods to do this are beyond the scope of this guide. Although it is possible to get good estimates of the phase equilibrium for 3-phase (gas-oilwater) systems, the available software does not allow rigorous simulation of threephase flow. The models present in Pipephase can only do two-phase (gas-liquid) flow calculations. Pipephase averages the properties of the liquid hydrocarbon and liquid water, and uses those average in the two-phase flow methods. Volumetric averaging, however, may not give good values for the viscosity and surface tension of the mixture. If the oil and water form an emulsion, the viscosity estimate may be off considerably using simple volumetric averaging, because the viscosity of an emulsion can be as much as 50 times as high as the viscosity of the oil or water. If it is likely that an emulsion will form, the Woeflin method, which is available in Pipephase, should be used to estimate the viscosity of the emulsion. 3.2.2

Pipeline Elevation Profile

The pipeline elevation profile used in the simulation can have a significant impact on the calculated pressure drop. Because the liquid holdup in upwardly inclined flow is greater than the holdup in downward flow, the elevational pressure drop in uphill legs is greater than the pressure recovery in downhill legs. As a result, elevational losses can account for much of the pressure drop in hilly terrain pipelines, even if the inlet and outlet of the line are at the same relative elevation. If the velocities in the line are high, the uphill and downhill holdups may be close. As the mixture velocity decreases, there will be an increasing difference between uphill and downhill holdups. The following table illustrates how sensitive the liquid holdup is to mixture velocity at various angles of inclination from horizontal. The feed stream is a gas-condensate with about 4 bbl/mm SCF of liquid present. (The values shown are predictions of the OLGAS model.)

CPTC

NOVEMBER 1994

31


MIXTURE VELOCITY FT/SEC ANGLE, DEGREES

2.7

4.1

5.4

8.1

16.2

-2.0

0.0041

0.0053

0.0064

0.0091

0.0115

-1.0

0.0052

0.0068

0.0085

0.0108

0.0122

-0.5

0.0068

0.0087

0.0108

0.0124

0.0126

0.0

0.0224

0.0218

0.0198

0.0156

0.0131

0.2

0.5797

0.4134

0.2249

0.0179

0.0134

0.5

0.5961

0.4988

0.3846

0.0317

0.0135

1.0

0.5997

0.5000

0.4314

0.3023

0.0144

2.0

0.6009

0.5024

0.4337

0.3428

0.0158

Using the values in the above table, a comparison of two models for a given section of a pipeline has been made. In the first model, the pipeline segment consists of two equal length sections of -0.5 degree and +0.5 degree each. The second model consists of a single horizontal pipeline segment. The liquid holdups for the two models are:

MIXTURE VELOCITY, FT/SEC

CPTC

HOLDUP FOR -0.5 DEGREE AND +0.5 MODEL

HOLDUP FOR HORIZONTAL MODEL

2.7

0.3015

0.0224

4.1

0.2538

0.0218

5.4

0.1977

0.0198

8.1

0.0221

0.0156

16.2

0.0131

0.0131

NOVEMBER 1994

32


The liquid holdups are far apart at low velocities and are the same at higher velocities. This comparison makes two points: • The pipeline profile must be realistic if the calculations of liquid holdup and pressure drop are to be accurate. • Low velocities cause severe problems in prediction of the pipeline performance. For very low velocities, it would be necessary to know the pipeline elevation profile within an accuracy of about one pipe diameter in order to get accurate holdup predictions. This is generally not practical. In many cases, the pipeline topography is not known when the preliminary pipeline sizing calculations are run. Frequently, in offshore pipeline design, the designer only knows water depths at subsea wells or platforms. Instead of assuming a straight line pipeline profile, it is recommended that the designer add some terrain features to the pipeline profile to simulate hills and valleys that are inevitably present in the actual profile. To improve the accuracy of the simulation, many calculation segments should be used in simulating the pipeline. Increasing the number of calculation segments always improves the accuracy of the simulation, but it increases the computer simulation time. The number of segments required depends on how rapidly the temperature, pressure and holdup are changing in the pipeline. For a system with rapid changes in pressure, e.g. flare systems, the number of calculation segments should be greater. If the temperature and pressure are changing slowly, a more coarse grid can be used to simulate the pipeline. 3.2.3

Heat Transfer

The temperature profile along the pipeline is important in many circumstances. The amount of condensation of liquids along a gas-condensate line, for instance, has a large impact on the pressure drop and liquid holdup in the line. Hydrate and wax deposition may occur in the line, requiring accurate estimates of temperatures. Corrosion is a strong function of temperature, so good heat transfer estimates are vital to corrosion prediction. To properly model the heat transfer between the pipeline and the surroundings, it is necessary to have information on the following:

CPTC

NOVEMBER 1994

33


•

thicknesses of the pipewall, pipeline coatings and insulation

•

whether the pipe is buried or exposed

•

the burial depth of the line

•

type of surroundings

•

ambient temperatures

•

thermal conductivities of the pipe, coatings and insulation.

With this information the programs can calculate heat transfer coefficients, which are then used to calculate the temperature profile in the pipeline. Values of the thermal properties for various materials can be read from the following table. Note that the Chevron Fluid Flow manual also has an extensive list of thermal conductivities for various types of materials.

Material

Thermal Conductivity, Btu/hr-ft-degF

Specific Heat, Btu/lb-degF

Density, lb/ft3

Carbon Steel

26

0.11

490

Stainless Steel

8-13

0.11

488

0.75-1.2

0.10

147-200

Onshore Soil (Wet)

1.35

0.20

90-110

Subsea Sandy Soil

1.25-1.50

0.30

105-115

Coal Tar Epoxy

0.20

0.35

92

Fusion Bonded Epoxy

0.15

0.32

75-90

0.12-0.15

0.50

90

0.011-0.022

0.38

2-12

Concrete (Saturated)

Neoprene Polyurethane Foam

CPTC

NOVEMBER 1994

34


At the early stages of a project, there may not be enough information to enable rigorous calculation of the heat transfer coefficient.. The following are rule of thumb values for heat transfer coefficients for subsea flowlines, which can be used in these instances:

U Value, BTU/hr/ft2/degF

Applications Wells

2

Risers

20-40

Buried Pipelines

1-3

Concrete Coated Nonburied Pipelines

3-5

Nonburied Pipelines without Concrete

5-10

For gas/condensate pipelines, temperature loss by the Joule-Thomson expansion (J-T) effect can be significant. In many gas pipelines, the temperature of the gas leaving the pipeline is less than ambient because of the J-T effect. Several concerns arise when using Pipephase for heat transfer calculations: a)

Pipephase only estimates temperature loss by the Joule-Thomson expansion cooling effect if the compositional model is used. The J-T effect is ignored in black oil simulations.

b)

The default velocity of water flowing past a pipeline is 10 miles per hour in Pipephase. This velocity is generally too high. More typical values are 1 to 3 ft/sec (0.7-2 mph).

c)

The Pipephase viscosity routine does not estimate viscosities at temperatures below 60 degrees F. At lower temperatures, it uses the viscosity at 60 degrees F. This can lead to errors for pipelines in deep water or cold climates.

d)

The thermal conductivity for saturated concrete is much higher than that for dry concrete. The saturated concrete value should be used for subsea pipelines with concrete coating.

CPTC

NOVEMBER 1994

35


e)

f)

Unless a value is entered for Hrad, radiation is ignored in the heat transfer calculations. For subsea or buried pipelines, radiation is negligible, but it can be a significant effect for surface flowlines. The convective heat transfer routines in Pipephase are not very rigorous. Errors in heat transfer calculations can occur for systems in which convection is the prime source of heat transfer.

3.2.4

Recommended Methods for Pressure Drop, Liquid Holdup, and Flow Regime Prediction

There have been hundreds of multiphase flow design methods developed in the past 50 years. Most computer programs contain dozens of options to select for pressure drop, liquid holdup, and flow regime predictions. Most of these methods only have small ranges in which their predictions are accurate. This section of the guide discusses this problem and gives some recommendations on which methods to use for certain applications. Most of the methods available in Pipephase are correlations based on data taken in small diameter (0.5-2 inch) test loops having an air-water flow operating at nearly atmospheric pressure. The correlations developed from these data sets frequently do not include the effects of all the key variables, such as pressure, because changes in these variables were not studied in the experimental work. These correlations extrapolate poorly from field conditions. In the past 10 years, the development of mechanistic modeling has created a marked improvement in prediction capabilities. As noted in Section 1.2, mechanistic models attempt to model the physical phenomena associated with each flow regime. Mechanistic models solve a set of simultaneous equations developed for a specific flow regime. Correlations for a few key parameters are required in order to solve the equation set. Mechanistic models extrapolate to field conditions much better than correlations because the mechanistic models account for the effects of all the major variables. Several mechanistic models have been developed in the past few years. Tulsa University has developed models for near vertical flow (Ansari) and a general model covering all inclinations (Xiao). The physics in these models are good, but the correlations built into them are based solely on small diameter, low pressure data. The OLGAS model is

CPTC

NOVEMBER 1994

36


currently the best available method for general multiphase flow calculations. OLGAS is based on a wide range of data (diameter from 1 to 8 inches, pressures from atmospheric to 1400 psi), and it extrapolates best to field conditions. OLGAS is a proprietary program that has not been available within Chevron. As this guide is being written, however, negotiations are underway to add OLGAS to Pipephase and several other programs as options. If OLGAS becomes available, it is the recommended method for prediction of pressure drop, liquid holdup and flow regime. Methods are available that are as good or slightly better than OLGAS in certain ranges, but they are not as good overall. The following methods can be used in Pipephase as a check of OLGAS or as the design method if OLGAS is not available: a)

b)

CPTC

Pressure Drop 1)

Near Horizontal Low Gas-Oil Ratio - Beggs and Bril1 (Moody) is good.

2)

Near Horizontal Gas/Condensate - Eaton-Oliemans is good for relatively high velocities. All of the methods are poor for low velocities.

3)

Near Vertical Gas/Condensate - Both Gray and Hagedorn-Brown are good.

4)

Near Vertical Gas/Oil - Hagedorn and Brown is good.

5)

Inclined Up - Nothing is good; Beggs and Brill (Moody) is fair.

6)

Inclined Down and Vertical Down - Everything is poor. Use Beggs and Brill (Moody), but answers may be suspect at times.

Liquid Holdup

_

1)

Near Horizontal Low Gas-Oil Ratio - Beggs and Brill (Moody) is O.K.

2)

Near Horizontal Gas/Condensate Lines - All available methods are poor. The Eaton holdup correlation is better than the other methods.

3)

Near Vertical Gas/Condensate - The most accurate method is no-slip.

4)

Near Vertical Gas/Oil - Hagedorn and Brown is pretty good.

5)

Inclined Up - Beggs and Brill (Moody) is usable for low GOR lines, nothing is accurate for gas/condensate. If gas velocities are high, use no-slip; otherwise use

NOVEMBER 1994

37


Beggs and Brill (Moody). The user must be careful because the holdups can be a factor of 10 in error in some cases. 6)

c)

Inclined Down and Vertical Down - Everything is poor. Use Beggs and Brill (Moody), but answers may be suspect.

Flow Regimes 1)

The Taitel-Dukler flow regime map is as good as OLGAS for near horizontal flow with the exception of the slug-dispersed bubble boundary. This boundary is very poorly predicted. If this method is used, it is recommended that a value of ~10 ft/sec be used as the superficial liquid velocity for the slug-dispersed bubble transition rather than the Taitel-Dukler prediction.

2)

The Taitel-Dukler-Barnea map for near vertical flow is also as accurate as OLGAS.

On occasion, the conditions for a simulation may cause otherwise good multiphase flow methods to give erroneous results. It is usually a good idea to spot-check the results by use of another method to ensure that the answers are reasonable. If there is a wide variance in results, CPTC should be contacted for guidance. 3.2.5 Interpretation of Results When a multiphase simulator such as Pipephase is run, the interpretation of the results can be difficult. The following section provides assistance in understanding Pipephase output, and ensuring that the design criteria for the line (velocities, flow regime, and allowable pressure drop) are met. As discussed in Section 2.2, the velocity in the pipeline should be kept within a limited range. Calculation of the velocities from a Pipephase output is not straightforward. The designers of Pipephase chose to include the actual gas and liquid velocities in their output table rather than the superficial gas and liquid velocities which are needed in the erosional velocity calculations. As discussed in Section 1.2, the superficial and actual velocities are related by simple formulas:

Vsg = Ug (1 − H l )

CPTC

NOVEMBER 1994

38


and

Vsl = U l H1

The liquid holdup is read from the "slip holdup" column. This calculation is made more difficult by the poor formatting of the liquid holdup in the Pipephase output table. (The liquid holdup is shown to only two decimal places in the table. For gas-condensate lines, if the liquid holdup is below 0.5 percent, the printout will show 0.00 for the holdup.) A more accurate way of calculating the superficial velocities from the Pipephase output tables which doesn’t rely on reading the value for the liquid holdup is:

Hl =

(U g − Vm) (U g − U l )

Vsl = U l H l Vsg = Vm − Vsl

To calculate the C value in the API-RP14E equation, the value of the no-slip mixture density must be known. Pipephase apparently only calculates and tabulates this value in the output table if the Beggs and Brill (Moody) method is used. If other methods are used, a value of 0.00 is given in the output table for the no-slip mixture density. The no-slip mixture density can be calculated, however, from the phase densities shown on the output table and the superficial velocities calculated above:

ρns =

(ρg Vsg ) + (ρl Vsl ) Vm

Pipephase allows the user to print a flow regime map based on either the Taitel-Dukler map for near horizontal flow or the Taitel-Dukler-Barnea map for near vertical flow. The flow regime map is printed only for the last "device" in a "link". If the "link" contains several pipes with different inclinations, the flow regime map for some of these sections may be quite different from the map at the last "device". The only way to print the flow regime map at specific points along the line is to make these points ends of Pipephase "links".

CPTC

NOVEMBER 1994

39


The "link" summary tables print the flow regime predictions for each pipeline segment. The printout shows both the predictions of the multiphase flow design method (e.g. Beggs and Brill) and the Taitel-Dukler method. If OLGAS is available, the flow regime predictions of OLGAS can be compared directly with the Taitel-Dukler prediction, and the user can feel confident that the predicted flow regime is valid if the two methods match. If methods other than OLGAS are used, disregard their flow regime predictions and only consider the Taitel-Dukler predictions as reasonable. Once the flow regime is determined, the designer needs to decide if this flow regime is acceptable. This decision is more difficult than it may appear. Ideally, the flow line should not be in the slug flow regime. In practice, it may be very difficult to design a line to avoid slug flow under all anticipated flow conditions. The only variables the designer can change are diameter and operating pressure; the changes in these variables required to avoid slug flow may be impractical. It should be pointed out that many pipelines operate successfully in slug flow. As long as the pipeline and downstream equipment are designed with proper consideration of slug flow effects, they can be successfully operated. The flow regime analysis may show that the line is in stratified flow. In many instances, this is an excellent flow regime in which to operate. At low flowrates, however, slugging may occur in lines predicted to be in stratified flow, induced by the terrain. Terrain induced slugs are generally much longer than the slugs in normal slug flow and can cause severe operating problems. Terrain slugging is discussed in more detail in Section 5.2.2. If the pressure drop and velocities for lines in dispersed bubble or annular flow are within acceptable limits, these flow regimes are usually good regimes in which to operate. The pressure drop in the line should be compared with the allowable pressure drop. The pressure drop in the line can be read from the Pipephase "link summary" table. It should be pointed out that pressure drop is not always a maximum at the highest flowrate. If the pipeline contains inclined or vertical elements, it is possible that the highest pressure drop may occur at a low flow condition due to high elevational losses at low flows. It is worthwhile to emphasize the point that the pipeline design should be checked at offdesign points as well as the nominal design point. For most pipelines, worst case conditions for liquid holdup and flow regime occur at turn-down conditions.

CPTC

NOVEMBER 1994

40


3.3

Section Highlights


Compositional models should be used for any gas-condensate or volatile oil system. This recommendation covers gas-oil ratios above 3500 SCF/bbl.

•

General practice with Pipephase: use the black oil model for lower gas-oil ratio streams.

•

If it is likely an emulsion will form, the Woeflin method (available in Pipephase) should be used to estimate the viscosity of the emulsion.

•

The pipeline profile must be realistic if the calculations of liquid holdup and pressure drop to be accurate.

•

Low velocities cause severe problems in prediction of the pipeline performance.

•

If OLGAS becomes available, it is the recommended method for prediction of pressure drop, liquid holdup, and flow regime.

•

Mechanistic models extrapolate to field conditions much better than correlations, since the mechanistic models account for the effects of all the major variables.

•

The following methods can be used in Pipephase as a check of OLGAS or as the design method if OLGAS is not available: 1.

Pressure Drop a) Near Horizontal Low Gas-Oil Ratio - Beggs and Brill (Moody) is good. b) Near Horizontal Gas/Condensate - Eaton-Oliemans is good for relatively high velocities. All of the models are poor for low velocities. c) Near Vertical Gas/Condensate - Both Gray and Hagedorn-Brown are good. d) Near Vertical Gas/Oil - Hagedorn and Brown is good. e) Inclined Up - Nothing is good; Beggs and Brill (Moody) is fair. f) Inclined Down and Vertical Down - Everything is poor. Use Beggs and Brill (Moody), but answers may be suspect at times.

CPTC

NOVEMBER 1994

41


Liquid Holdup g) Near Horizontal Low Gas-Oil Ratio - Beggs and Brill (Moody) is O.K. h) Near Horizontal Gas/Condensate Lines - Nothing is accurate. The Eaton holdup correlation is poor, but better than the other methods. i) Near Vertical Gas/Condensate - The most accurate method is no-slip. j) Near Vertical Gas/Oil - Hagedorn and Brown is pretty good. k) Inclined Up - Beggs and Brill (Moody) issuables for low GOR lines, nothing is accurate for gas/condensate. If gas velocities are high, use noslip; otherwise use Beggs and Brill (Moody). Be careful because the holdups can be a factor of 10 in error in some cases. l) Inclined Down and Vertical Down - Everything is poor. Use Beggs and Brill (Moody), but answers may be suspect. Flow Regimes a) The Taitel-Dukler flow regime map is as good as OLGAS for near horizontal flow with the exception of the slug-dispersed bubble boundary. This boundary is very poorly predicted. If this method is used, it is recommended that a value of ~10 ft/sec be used as the superficial liquid velocity for the slug-dispersed bubble transition rather than the TaitelDukler prediction. b) The Taitel-Dukler-Barnea map for near vertical flow is also as accurate as OLGAS. •

The flow line should, ideally, not be in the slug flow regime. In practice, it may be very difficult to design a line to avoid slug flow under all anticipated flow conditions.

•

At low flow rates slugging may occur in lines predicted to be in stratified flow, induced by the terrain.

•

If the pressure drop and velocity for lines in dispersed bubble or annular flow are within acceptable limits, these flow regimes are usually good regimes in which to operate.

CPTC

NOVEMBER 1994

42


SECTION 4.0 - TRANSIENT FLOW MODELING 4.1

Transient Flow Modeling (General)

Transient multiphase flow simulators have only been developed recently. The first widely used commercial program, OLGA, began development in about 1983 and has been commercially available since 1990. OLGA’s only current competitor, PLAC, was introduced to the market at about the same time. Chevron currently does not own either program but has used OLGA for specific projects through consultants. Chevron internally developed a transient code, Transpire, in the same time frame as OLGA. This program has not been widely used, and it does not have as many features as the commercial codes. Steady state simulators assume that all flowrates, pressures, temperatures, etc. are constant through time. Inherently transient phenomena, such as slug flow, are modeled by use of their average holdups and pressure drops. Transient models allow all the input variables to change with time. Transient programs can model phenomena such as slug flow and can show the variations in parameters such as outlet gas and liquid flowrates as a function of time. Transient simulators, therefore, model the actual operation of pipelines closer and with more detail than steady state simulators. Transient simulators solve a set of equations for conservation of mass, momentum and energy to calculate the liquid and gas flowrates, pressures, temperatures and liquid holdups. These calculations are done at each time step. The programs utilize an iterative procedure, which ensures that a set of boundary conditions (such as inlet flowrates and outlet pressures as a function of time) are met while solving the conservation equations. Steady state modeling can be used to size pipelines, but the predicted size may be inaccurate if the line is in slug flow. Transient simulators can size pipelines more accurately, and they are valuable in several other areas such as the design of downstream facilities, development of operating guidelines, and the diagnosis of operating problems. Steady state simulators cannot properly address any of these other concerns. 4.2

Use of Transient Simulators

Because of their power, transient simulators have been used for a variety of purposes. These uses include: a)

CPTC

Slug flow modeling

NOVEMBER 1994

43


b)

Estimates of the potential for terrain slugging

c)

Pigging simulation

d)

Estimation of corrosion potential in low spots in the line

e)

Startup, shutdown and pipeline depressuring simulations

f)

Development of operating guidelines

g)

Real time modeling, including leak detection

h)

Operator training

i)

Design of control systems for downstream equipment

j)

Slug catcher design

A general guideline for the use of steady state and transient modeling would be to use steady state modeling during the feasibility level design of a system but use transient modeling in the final design of the pipeline and its associated equipment. As transient simulators improve and computer power increases, it is likely that transient simulators will eventually supplant steady state simulators. Because Chevron does not own a transient simulator at this time, this guide does not contain any guidelines for their use. Section 5.1 discusses the use of the OLGA program for slug length prediction.

4.3

Section Highlights


Transient simulators model the actual operation of pipelines much closer than steady state simulators.

•

General guideline for the use of steady state and transient modeling: use steady state modeling during the feasibility level design of a system, but use transient modeling in the final design of the pipeline and its associated equipment.

CPTC

NOVEMBER 1994

44


SECTION 5.0 - SLUG FLOW ANALYSIS 5.1

Slug Flow - General

The formation of slugs of liquid can be caused by a variety of mechanisms: a) Hydrodynamic Slugging b) Terrain Slugging c) Pigging d) Startup and Blowdown e) Flowrate Changes Each of the mechanisms will be briefly discussed here, and will be further discussed in Section 5.2. Hydrodynamic slugging refers to operating in the slug flow regime. In near horizontal flow, slugs are formed by waves growing on the liquid surface to a height sufficient to completely fill the pipe. When this happens, alternating slugs of liquid and bubbles of gas flow through the pipe, as illustrated in Figure I:1-1. Terrain slugging occurs when a low point in the line fills with liquid. The liquid does not move until gas pressure behind the blockage builds high enough to push the liquid out of the low spot as a slug. Terrain slugging can produce very long slugs in pipeline-riser systems. Although terrain slugging occurs at low superficial gas and liquid velocities, the actual velocities during slug release can be very high. When a pipeline is pigged, most of the liquid inventory is pushed from the line as a liquid slug ahead of the pig. When a line is shut down, liquid that is left in the line will drain to the low points in the line. When the flow is restarted, the accumulated liquid may exit the pipeline as a slug.

CPTC

NOVEMBER 1994

45


When the flowrate is increased, the liquid holdup in the line decreases. This change in holdup can either exit the line as a gradual increase in liquid flow, or it can come out in the form of a slug, depending on the flowrate. Each of the slug flow mechanisms is highly transient in nature. Steady state models cannot properly simulate slug flow behavior and are very limited in their ability to predict slug characteristics such as slug length and frequency. The next two sections of the guide discuss the slug flow mechanisms in more detail, discuss available methods of predicting slug flow behavior and give some recommendations on sizing of slug catchers and separators. 5.2

Slug Length and Frequency Predictions

Although estimates of slug length and frequency are of prime importance in design of pipeline system facilities, most of the prediction methods available are poor. Development of prediction methods has been hampered by the difficulty of the problem and the meager amount of available test data. This section discusses each of the mechanisms for slug flow, discusses the available test data, and give recommendations on the best available prediction methods. 5.2.1

Hydrodynamic Slugging

Experimental measurements of the slug length in hydrodynamic slug flow show several interesting results: a)

The slug length is not constant. At a given point in the line, the slug length varies greatly around an average value. Different investigators have characterized the slug length distribution as log normal, truncated Gaussian, inverse Gaussian, or fractal distributions. The maximum slug length may be several times greater than the average.

b)

The average slug length and the slug length distribution change with the position down the pipe. Slugs may grow, dissipate, or merge as the flow continues down the pipe. As a result, the average slug length usually increases with the position in the pipe, while the standard deviation of the slug length distribution decreases.

CPTC

NOVEMBER 1994

46


c)

Slugs in vertical pipes are much smaller than slugs in horizontal pipes.

d)

The slug length in laboratory experiments can be fairly well correlated. These tests show that the average slug length (in feet) is approximately 32 times the Pipe Diameter (in feet) for horizontal pipes.

e)

In the data base of published pipeline field test results, the average slug length is much higher than the results observed in the laboratory. The field tests results show average slug lengths of 300-2000 times the Pipe Diameter, with some slugs as long as 10,000 times the Pipe Diameter.

The differences between laboratory and field data shown in points d) and e) above are due to factors such as: -

terrain features have a large effect on the slug length and frequency;

-

slug flow in the field can be combination of mechanisms such as hydrodynamic slugging causing terrain slugging;

-

field pipelines are much longer, allowing more time for slug growth.

Average slug length is a complex function of many variables: the diameter and length of the pipeline; the topography of the line; the gas and liquid superficial velocities; the liquid physical properties; and the gas density. Several correlations have been presented for the prediction of slug length and slug frequency for horizontal piping and pipelines. Most of these correlations are based solely on laboratory data, which means they are of limited use in the design of pipelines in the field. A few correlation methods have been presented based on field data. Two of these methods, the Brill, et al. Correlation and the Hill & Wood method, have been widely used for slug length prediction. Both methods will be discussed in detail. Brill, et al. took several sets of data on 12 and 16 inch pipelines at Prudhoe Bay in about 1978. They were the first experimentalists to report the wide disparity between the extrapolation of lab results and field data. They developed a simple correlation for slug

CPTC

NOVEMBER 1994

47


length based on the Prudhoe Bay tests and small diameter laboratory data. Their correlation is: ln(Ls) = -2.663 + 5.44 (ln(D))0.5 + 0.059 ln(Vm) where Ls

= Average slug length, ft.

D

= Pipeline inside diameter, inches

Vm

= Mixture velocity, ft/sec

The Prudhoe Bay test data appeared to be a log normal distribution around the mean slug length. Log normal distributions were fit to each of the Prudhoe Bay tests, and the mean slug length and variance were calculated. Their method assumes that the variance for any pipeline is the same as the average value of the variance from these tests. With this assumption, the slug length distribution is the same for all pipelines. The distribution is:

Percent Probability

Slug Length/Mean Slug Length

50.0000

1.00

84.1300

1.65

97.7200

2.72

99.8600

4.46

99.9900

6.42

99.9999

10.76

The Brill method is easy to use, and it has been used as the design basis for many facilities. The Brill et al. model is the basis for the slug length predictions in Pipephase. Unfortunately, the Brill method is very inaccurate. It gives poor predictions for almost every data set that was not included in the original correlation. A comparison of the Brill predictions against measured average slug lengths from a 16" laboratory line and two field pipelines shows:

CPTC

NOVEMBER 1994

48


Actual

Predicted

Avg. Slug Length, ft.

Avg. Slug Length, ft.

30

700

Field Data

1000

120

Field Data

400

1800

Data Source BHRG

The inaccuracy of the Brill et al. method is due to the simplicity of their formula. The slug length is predicted to be almost exclusively a function of the line diameter. As noted previously, the slug length is a complex function of many variables, and a simple formula like this cannot approximate reality. There are two other methods based primarily on the Prudhoe Bay data, namely, the Norris correlation and the Scott, et al. correlation. Their performance, while a bit better than the Brill method for larger pipelines, is also weak. Both of these methods assume that the slug length is only dependent on the line diameter. In 1990, Hill and Wood of BP published a paper proposing an alternate way of modeling slug frequency. Their work was based on both laboratory data and a large number of field measurements. The model is more sophisticated than the Brill approach and attempts to account for the many of the major variables. Their model correlates the slug frequency with the diameter, gas-liquid slip velocity, and the equilibrium holdup at the beginning of the pipeline. The model assumes a horizontal pipe and uses the Taitel-Dukler stratified flow model to estimate the slip velocity and the equilibrium holdup at the beginning of the pipeline. In order to calculate the Taitel-Dukler holdup, the value for the Lockhart & Martinelli X factor is calculated from the following equation:

V  X =  sl   Vsg 

where:

0.9

 ρ1    ρ   g

0.4

 µ1    µ   g

0.1

µ 1 = Liquid Viscosity, cp µ g = Gas Viscosity, cp

CPTC

NOVEMBER 1994

49


Once X is known, the value for the liquid holdup (Hle) for stratified flow can be read from Figure I:5-1. The actual gas and liquid velocities can be calculated from the superficial velocities and the stratified flow liquid holdup by the formulas: Ug =

Vsg 1 − H le

Ul =

and

Vsl H le

Their equation for slug frequency is:

Fs =

(

2.74 H le U g − U l D(1 − H le )

where Fs

)

= Slug Frequency, slugs/hour

D

= Pipe Inside Diameter, ft

Ug

= Actual Gas Velocity, ft/sec

Ul

= Actual Liquid Velocity, ft/sec

Hle

= Liquid Holdup (based on stratified flow at the inlet of the pipe)

To calculate the slug length, the slug fraction, which is defined as the slug length divided by the sum of the slug and bubble lengths, must be known. The slug fraction can be calculated rigorously by use of mechanistic slug flow models, such as the Hubbard-Dukler model or the Nicholson, Gregory, and Aziz model. A simplified alternative to the use of these models is the following procedure. First, the liquid holdup in the slug will be estimated. The Gregory, Nicholson, and Aziz method is the easiest method to use, and it is as accurate as most of the methods. The liquid holdup in the slug is given by:

CPTC

NOVEMBER 1994

50


H ls =

where Hls

1 1.39 V  1+  m   28.4 

= Liquid Holdup in the slug

The liquid holdup in the bubble (Hlb) can be assumed to be ~0.20 for much of the slug flow range. From a material balance, the slug fraction (SF) can be calculated from these values and the overall liquid holdup prediction:

SF =

Hl − H lb Hls − H lb

The slug length can be calculated from:

Ls = where Ls

3600SF (Vm ) Fs = Mean Slug Length, ft

An example of the use of the Hill & Wood model is shown in Section 6.1. To use the Hill & Wood model for slug catcher designs, the design slug length is needed instead of the mean slug length. A rule of thumb, based on the limited amount of experimental data available, is that the maximum slug length is approximately 6 times the mean slug length. The Hill & Wood model is more sophisticated than the Brill model and is based on a broader data base. If a correlation based model is used for slug length prediction, the Hill & Wood model is probably the best available method. Another alternative approach to hydrodynamic slug length prediction is the method used in the OLGA transient program. The slug tracking model in OLGA attempts to dynamically track each slug in the system, from its inception to its dissipation or exit from the pipeline. This method models the growth or dissipation of each slug and models phenomena such as the merger of two slugs. The model accounts for terrain effects and

CPTC

NOVEMBER 1994

51


the effects of all of the major variables. This model, which has only been available to the industry since early 1994, shows promise in prediction of slug lengths. 5.2.2 Terrain Slugging Terrain slugging refers to slug flow resulting from a blockage of a low spot in the pipe with liquid. When the liquid completely fills up a low spot, the liquid remains stationary until gas pressure behind the slug builds to a suitably high level to push the liquid out as a slug. The most severe form of terrain slugging occurs in pipeline-riser systems, which have a negatively sloping pipeline ahead of the riser. At low flowrates very large slugs may form in this type of system. These slugs can be much larger than those associated with hydrodynamic slugging. Figure I:5-2 illustrates the various stages in terrain slugging for pipeline-riser systems. Experimental work has shown that both hydrodynamic and terrain slugging may occur at various points in the line, under certain conditions. This is one of the reasons that it is so difficult to extrapolate laboratory slug length data to field conditions. One of the problems with using a steady state program such as Pipephase is its inability to identify terrain slugging. Pipelines in the terrain slugging regime will usually be identified by the steady state simulator as being in stratified flow. If no further analysis is done, the line may be designed in the belief that it is in a steady flow regime. On startup, the line may experience slugs that may be thousands of feet long. The best way to analyze the line for terrain slugging is to run a transient simulator. If transient analysis is not possible, some simplified methods have been developed for estimation of severe slug potential. These methods pertain only to the modeling of pipeline-riser slugging. The method of Pots et al. Has been included in Pipephase, and it is discussed here. By equating pressure buildup rates, Pots developed a formula for a dimensionless severe slugging factor πss:

CPTC

NOVEMBER 1994

52


 144  zRT   Wg   gc      π ss =    M   Wl   g   L f (1 − H l ) where z

= Gas compressibility

R

= Gas Constant, 10.73 psia-ft3/(lb-mole-deg R)

T

= Temperature, deg R

M

= Gas Molecular Weight, lb/lb-mole

Wg

= Gas mass flowrate, lb/sec

Wl

= Liquid mass flowrate, lb/sec

gc

= Conversion factor, 32.2 lb-ft/lbf-sec2)

g

= Acceleration due to gravity, 32.2 ft/sec2

Lf

= Length of the pipeline, ft

Hl

= Liquid holdup in the pipeline

When the value for πss is less than 1, severe slugging could occur. Values less than 1 do not ensure that severe slugging will occur. Pots indicates that the flowline must be in stratified flow for severe slugging to take place. The Taitel-Dukler flow regime map can be used to determine whether the flowline is in stratified flow at the pipeline-riser junction. Pots also gives an expression for the slug length in severe slugging: Ls = Lr / πss where Ls

= Slug length, ft

Lr

= Riser height, ft

Pipephase calculates πss if the SLUG table option is turned on. The equation for πss indicates that long pipelines, low pressure, and low gas-oil ratios contribute to the likelihood and severity of terrain slugging.

CPTC

NOVEMBER 1994

53


The Pots method gives a reasonable approximation to terrain slugging behavior for a simple constant slope pipeline-riser system, but it cannot rigorously model the performance of a hilly terrain pipeline connected to a riser. 5.2.3 Pigging Slugs Pigs are run through pipelines for a variety of reasons, including: a) Liquid inventory control b) Maintenance and data logging c) Pipeline cleaning and dewaxing d) Inhibitor application The classic analysis of pigging was developed by Baker and McDonald in 1959. They identified five zones in a pipeline undergoing pigging: a) Undisturbed equilibrium flow b) Liquid slug ahead of the pig c) Pig itself d) Dry gas region behind the pig e) Reestablished equilibrium flow Figure I:5-3 graphically illustrates the zones in a pigged pipeline. Baker and McDonald developed a quasi-transient model of pigging in which the lengths and pressure drops for the various zones were estimated and tracked with time. Although some of the methods used in the Baker-McDonald model are antiquated, the model is conceptually a very good representation of pigging. Several investigators have modified the Baker-McDonald model to improve its predictions. The model developed by Barua of Tulsa University has been included in Pipephase.

CPTC

NOVEMBER 1994

54


The commercial transient simulators model the physics of pigging more rigorously than the models of Baker & McDonald or Barua, and are more accurate. The simpler models, however, give pretty good predictions. It is also possible to get a good estimate of the pig volume and transit time by simple hand calculations. To estimate the slug volume and transit time by hand, the designer can use the following method. The pig and the slug in front of it move at a velocity that is equal to the gas velocity behind the pig. For a first approximation, this velocity can be assumed to be equal to the mixture velocity. The transit time of the pig would therefore be: ttrans where ttrans Lf

= Lf / Vm avg = Transit time for the pig, sec = Pipeline length, ft

Vm avg = Average mixture velocity, ft/sec When the slug exits the line, the volume of liquid ahead of the pig is equal to the liquid holdup in the line minus the amount of leakage past the pig minus the amount of liquid produced while the pig is traversing the pipeline or: Qslug where Qslug

= (Hl Ap Lf) (1 - fleak) - Vsl Ap ttrans = Volume of liquid ahead of the pig, ft3

Hl

= Average liquid holdup in the pipe

Ap

= Cross-sectional area of the pipe, ft2

fleak

= Fraction of the liquid that leaks past the pig

The value for fleak is dependent on the type of pig and the pig velocity. In lieu of better data, use fleak = 0.02. The amount of time during which the liquid slug enters the slug catcher is:

CPTC

NOVEMBER 1994

55


t slug =

where tslug

Qslug A p Vm avg = Time during which the slug enters the slug catcher, sec

The liquid rate while the slug enters the slug catcher is: qslug where qslug

= Vm avg Ap = Liquid rate into slug catcher when slug is exiting, ft3/sec

The simplified analysis shown above doesn’t account for many effects that occur during pigging, such as pig acceleration and deceleration, aeration of the liquid slug, etc., but it gives a ball-park estimate of the flows and volumes needed for the slug catcher design. 5.2.4

Startup and Blowdown Slugs

When a pipeline is shut down, the liquid will drain to the low points in the line. When the line is restarted, this liquid may exit the line in the form of slugs. To determine whether slugging will occur, and to estimate the magnitude of the slugs, a transient simulator must be used. If the pipeline is depressured at shutdown, slugs may also form, due to high gas velocities during the blowdown period. Transient simulation is needed to model these slugs. 5.2.5

Rate Change Slugs

When the flowrate is increased, the liquid holdup in the line decreases. This change in holdup can either exit the line as a steady flow with increased liquid production, or it can come out in the form of a slug, depending on the flowrate change. Please note that the rate change slugs can occur in gas/condensate lines when the rates are increased. The line may be in a steady flow pattern, such as stratified flow, at both the initial and final flowrates but will slug during the transition period until the line reequilibrates at the higher rate.

CPTC

NOVEMBER 1994

56


As with startup slugs, it is impossible to predict whether slugs will occur when rates are changed using steady state or hand methods. The line must be dynamically simulated using a transient flow program. 5.2.6

Downstream Equipment Design for Slug Flow

The design of slug-catchers, separators, and control systems downstream of pipelines must comprehend the presence and severity of slug flow. Estimates of parameters such as slug volumes, liquid and gas rates exiting the pipeline as a function of time, etc. must be factored into the design of this equipment. These variables should be calculated for the design operation and for a series of off-design cases: turndown rates, pigging, shutdown and startup, rate changes, etc. As mentioned, transient modeling gives the best estimates of slug flow behavior. Some of the transient simulators also allow the user to simulate separators and control systems as part of the run, allowing the user to fine tune the design. For approximate sizing of the equipment, the methods shown in Sections 5.2.1 to 5.2.3 can be used. Off-design cases should be estimated as well as the design point. In general, the size of slug catching equipment for gas-condensate pipelines will be governed by pigging considerations. For oil dominated systems, the size of the slug catcher is usually governed by the maximum slug length due to either hydrodynamic or terrain slugs. Because slug catching equipment can be a substantial cost item, it is possible to minimize the cost of the equipment by considering alternative operating scenarios. Some of these include: a)

CPTC

Does the line need to be pigged routinely or is pigging only needed for maintenance? If pigging is a maintenance item done once per year, it is possible to run the pig at a rate low enough to keep a small slug catcher from overflowing. It is also possible to send the liquid from the slug catcher to low pressure separation at high rates for short periods, if the occurrence of pigging is infrequent. This, too, minimizes the high pressure slug catcher sizing.

NOVEMBER 1994

57


b)

Pigging of the line after steady state conditions are reached may result in a very large slug catcher. If a high pigging frequency is chosen, so that the line never achieves steady state operation, the slug catcher size can be minimized.

c)

It is possible to design the line for frequent pigging at low rates and steady state operation at higher rates, thereby minimizing slug catcher size while eliminating the need for frequent pigging when the rate is high.

d)

If pigging isn’t necessary for a gas-condensate line, the sizing of the slug catcher usually becomes dependent on rate changes. Limiting the turndown on the line and limiting the amount that the rate is changed at one time can be beneficial in minimizing slug catcher sizes.

e)

The use of parallel pipelines should be considered as a way of limiting the turndown for a line. When rates are low, all the production can be fed to one of the lines, thereby keeping velocities high, minimizing liquid holdup.

f)

If terrain slugging is shown to be a problem, there are several remedial steps which have been employed to decrease the severity of the slugging or eliminate it completely. One approach would be to reroute the pipeline to reduce or eliminate dips in the pipeline profile. Two other methods are: choking of the flow at the tope of the riser; and gas lifting the riser.

g)

Because the slug length for hydrodynamic slug flow is a function of the diameter of the line, the use of parallel pipelines instead of one pipeline can decrease the slug length and volume going to the slug catcher/separator.

h)

The use of multiphase pumps or subsea separators at well clusters can decrease slug catcher sizes significantly.

i)

Pigs have been developed with a variety of proprietary internals that limit the velocity at which the pig moves through the pipeline. It may be possible to limit the slug catcher size by restricting the pig velocity, which, in turn, limits the rate of liquid exiting the line during the pigging operation.

CPTC

NOVEMBER 1994

58


5.3

Section Highlights


The formation of slugs of liquid can be caused by a variety of mechanisms: 1. Hydrodynamic Slugging 2. Terrain Slugging 3. Pigging 4. Startup 5. Rate Changes

•

The Brill method is very inaccurate. It gives poor predictions for almost every data set that was not included in the original correlation.

•

Rule of thumb (based on limited amount of data available): the maximum slug length is approximately 6 times the mean slug.

•

If a correlation based model is used for slug length prediction, the Hill & Wood model is probably the best available method.

•

Pipelines, which are in the terrain slugging regime, will usually be identified by the steady state simulator as being in stratified flow. If no further analysis is done, the line may be designed in the belief that it is in a steady flow regime. On startup, the line may experience slugs that may be thousands of feet long..

•

In general, the size of slugcatching equipment for gas-condensate pipelines will be governed by pigging considerations. For oil dominated systems, the size of the slugcatcher is usually governed by maximum slug length due to either hydrodynamic or terrain slugs.

•

There are alternative operating scenarios which can be considered to minimize slug catcher size.

CPTC

NOVEMBER 1994

59


Figure I: 5-1 Taital-Dukler Liquid Holdup Predictions for Simplified Flow in a Horizontal Pipeline

CPTC

NOVEMBER 1994

60


Figure I: 5-2 Stages in Terrain Slugging of Pipeline Riser Systems

CPTC

NOVEMBER 1994

61


Figure I: 5-3 Pipeline Pigging

CPTC

NOVEMBER 1994

62


SECTION 6.0 - EXAMPLE PROBLEMS 6.1 EXAMPLE1: Low Gas/Oil Line Between Platforms This example is based on the Mejo to Tapa P/P line in Nigeria. The line runs 8 miles from a Wellhead Platform to a Producing Platform. The gas-oil ratio for the production is about 345 SCF/bbl oil. Water production is initially zero, but builds rapidly with time. The production profile is: Year No.

Oil Production, bbl/day

Water Production, bbl/day

Gas Production, MMSCFD

1

15,478

0

5.32

4

7,954

8,865

2.74

8

2,057

7,584

0.70

12

765

4,438

0.26

The API gravity of the stock tank oil is 27.0. The specific gravity of the separator gas is about 0.71. The dry wellstream composition, as measured by a separator test, is: Component CO2 N2 C1 C2 C3 iC4 nC4 iC5 nC5 C6’s C7’s C8’s C9’s C10’s

CPTC

Mole % 0.90 0.05 31.94 4.18 1.55 0.45 0.49 0.48 0.37 1.66 2.04 4.23 3.55 3.45

NOVEMBER 1994

63


C11’s C12's Cl3’s Cl4’s C15's C16's Cl7's C18's Cl9's C20+ Total

3.07 2.57 2.64 2.22 2.06 1.82 1.48 1.41 1.09 26.30 100.00

The specific gravity of the C20+ fraction is 0.9499 at 60 degrees F. and the molecular weight of the C20+ is 309. The hydrocarbon wellstream composition is assumed to remain constant throughout the life of the field. The pipeline system consists of a downward riser at the wellhead platform, the 8 mile long pipeline, and an upward riser at the production platform. The down riser is vertical and it is 65 feet long. The up riser is vertical and 100 feet long. The pipeline elevation profile is not known. The only information available on the pipeline elevation is that the water depth at the wellhead platform is 15 feet, and the water depth at the production platform is 20 feet. The pressure at the production platform is 80 psig. The maximum pressure at the wellhead platform is 300 psig. The line will be unburied, and it will not be coated with concrete. The wellstream temperature at the wellhead platform is 130 degrees F, and the ambient air and water temperatures are 80 degrees F. What line size is required? Is slugging likely to be a problem? If slugging occurs, what length and frequency are the slugs?

CPTC

NOVEMBER 1994

64


Solution The operation of the line will be initially simulated by Pipephase. If it appears that slugging is likely, additional simulations will be done using the OLGA transient model. 6.1.1

Line Size

In order to run Pipephase, an estimate of the line diameter is needed. The minimum line size is given by the erosional velocity formula, Vmax =

where

C ρ ns

C = 100

To calculate ρns and Vmax rigorously, the phase equilibria and physical properties would need to be found by flash calculations. Since the minimum diameter is only a starting place for the pipe sizing calculations, it isn’t necessary to calculate these values with great precision. When the Pipephase results are analyzed, the values of the erosional velocity constant will be computed and compared against the maximum value of 100. The maximum value of C should occur at the outlet of the up riser, since the mixture velocity is the highest there. The pressure at the top of the riser should be almost equal to the separator pressure, 80 psig. The temperature of the fluid at the top of the riser should be close to sea temperature, 80 degrees F. At these conditions, almost all the gas should be flashed out of solution, so that the amount of gas present should be equal to the gas/oil ratio times the oil production rate. The compressibility factor of the gas at these conditions would be about 0.98. Therefore the amount of gas and liquid present under the year 1 conditions should be approximately: Gas ft3/sec: cfsg = (345 SCF/bbl) (15478 bbL/D) (14.7 psia/94.7 psia) [(460+80) F/(460+60) F] (1 D/86,400 sec) .98 = 9.76 ft3/sec Liquid ft3/sec:

CPTC

NOVEMBER 1994

65


cfsl = (15478 bbl/D) (1 D/86,400 sec) (42 gal/bbl) (1 ft 3/7.48 gal) = 1.01 ft3/sec The densities of each phase are approximately: Gas:

ρg =

(.71) (28.97 lb/mole) (95 psia) / [(.98) (10.73 psia-ft3/mole-degF) (460+80 degF)] = 0.344 lb/ft3

Liquid:

ρl =

(62.4 lb/ft3) [141.5/(131.5+27 deg API)] = 55.7 lb/ft3

The no-slip average density of the mixture, ρns, is given by:

ρns =

(ρg cfsg ) + (ρl cfsl ) (cfsg + cfsl )

= [(0.344 * 9.76) + (55.7 *1.01)] / (9.76 + 1.01) = 5.54 lb/ft 3 The erosion velocity criteria gives a maximum velocity of: Vmax = 100 / (5.54)0.5 = 42.5 ft/sec The minimum diameter of the line, therefore, is: Dmin = [( 4 * (9.76 + 1.01) ft3/sec) / (42.5 ft/sec) * 3.1416)]0.5 (12 in/ft) = 6.82 inches Rounding this result to the nearest standard line size indicates that the pipeline must be at least an 8 inch line to meet the erosional velocity criterion. Pipephase runs should be made for the production rates at the various years starting with a line size of 8 inches. If the pressure drop becomes excessive, increase the line size, until

CPTC

NOVEMBER 1994

66


the inlet pressure for the limiting case remains below about 300 psig for all production rates. The approach discussed in the previous paragraph requires a trial and error solution for the diameter. To cut down the number of trial diameters, it is possible to approximate the frictional pressure drop by use of a "ball-park" method. The friction loss can be estimated by assuming a homogenous mixture of the two phases. The assumption of homogenous flow gives a friction loss that is low, typically being anywhere from 5% to 100% too low. Adding 25% to the homogenous friction loss will bring the answer into a reasonable range. For this example, the elevational and accelerational pressure drops will be much less than the frictional loss, and will be ignored in the rough line sizing calculations. Using typical values for the viscosities of each phase (Gas Visc. = 0.015 cp, Liq. Visc. = 10 cp) gives a no-slip viscosity of: [(0.015 x 9.76) + (10 x 1.01)] /(9.76 + 1.01) = 0.95 cp The homogenous frictional pressure gradient is: 2 dp 4 f nsV m ρns = 2 g c (12d ) dx

d 1488 ρnsVm    12  Rens = µns

Vm =

(9.76 + 1.01) ft 3 / sec x 4 x 144 3.1416 d 2

where d is in inches Vm = 1975 / d2 (ft/sec) ρns = 5.54 lb/ft3

CPTC

NOVEMBER 1994

67


µns = 0.95 cp Substituting these values into the expression for Rens gives: Rens = 1,426,000/d For a smooth pipe, the Fanning friction factor can be approximated by: fns = 0.046 / Rens0.2 Substituting, fns = 0.046/(1426000/d)0.2 = .00270 d0.2 Combining this expression with the expressions above gives the following expression for the homogenous frictional pressure gradient:

dp = dx

(

)

 1975  4 0.00270d 0.2  2    d (64.4)(12d)

2

(5.54)

= 302.0/d4.8 psi/ft The allowable pressure drop is 300-80=220 psi in the 8 mile long pipeline. Equating this pressure drop to the homogenous pressure gradient, and adding 25% to the homogenous pressure gradient to account for the two-phase behavior gives: 220 psi/(8 * 5280 ft) = 1.25 * 302.0/d4 8 psi/ft d = 10.29 inches These calculations indicate that an 8 inch line is probably too small, and that a 10 inch line is probably the minimum line size that can be used to stay within the desired pressure drop.

CPTC

NOVEMBER 1994

68


For the sake of this example, we will make Pipephase runs for 8, 10, and 12 inch pipelines and risers. Since we have both black oil properties and compositions available for this example, we will run both the black oil model and the compositional model in Pipephase for the 10 inch line as a comparison between the methods. In general, compositional modeling predicts the phase equilibrium more accurately than black oil modeling, but the black oil models may predict the physical properties of the liquid better than compositional modeling. For low GOR systems, black oil modeling may give the best answers, while compositional modeling is clearly superior for high GOR systems. The black oil model only needs the gravities of each phase and the total gas-oil ratio of the system. For the compositional model, the program needs two of the following for each petroleum cut: molecular weight; specific gravity; or normal boiling point. In the example information, we are given the specific gravity and molecular weight for the C20 + cut, but we have no information on the C6 to Cl9 cuts. Without this information, the only possible characterization is to use the normal paraffin data for the cut. Use nCl0 as the basis for the properties of the Cl0 cut, for instance. In reality, each of these cuts consists of hundreds of different components boiling in the same boiling range as the normal paraffin, and the phase behavior of this mixture of aromatics, naphthenes, and alkylated paraffins is not the same as the behavior of the normal paraffin. The proper way to characterize the cuts is to have laboratory analyses of the specific gravity and molecular weight of each of the cuts in addition to their boiling ranges, but this information is rarely available. Information is needed on the pipeline and risers. The diameters to be used in this exercise are: Nominal Diameter, inches

CPTC

Outside Diameter, inches

Inside Diameter, inches

8

8.625

7.625

10

10.750

9.500

12

12.750

11.375

NOVEMBER 1994

69


Since there is insufficient data to do rigorous heat transfer calculations, use an overall heat transfer coefficient, U, of 5 BTU/hr/ft2/degF for the pipeline and a U of 20 for each of the risers. The pipeline elevation profile has not been given. The simplest way to simulate the pipeline profile is to assume that it has a constant slope, with a 5 foot drop in elevation over its 8 mile length. This assumption will give the most optimistic results, because it will minimize the elevational pressure drop for hills. To better mirror reality, some terrain effects could be added by adding hills and valleys to the profile. For high velocities, there will be negligible difference in the holdup and pressure drop between the constant slope profile and the profile with the hills added. For low velocities, however, the difference can be significant. For the sake of simplicity in this example, we will use the straight-line profile for the Pipephase runs, with the understanding that the pressure drop and liquid holdup for low rates may be low. The Pipephase runs will be made with a constant outlet ("Sink") pressure of 80 psig. The inlet ("Source") pressure to the line will vary with the production rate. Since this line is a low gas-oil ratio line, the Beggs and Brill correlation will be used to calculate the pressure drop and liquid holdup for the line. If the OLGAS method had been available, it would have been used for the calculations. Because of its mechanistic model, OLGAS would show more of an impact of hills than Beggs and Brill, and would give more accurate answers. The Beggs and Brill flow regime calculations are poor, and they will be ignored. The Taitel-Dukler-Barnea flow regime predictions will be used instead. A comparison of the predictions of the black oil model with three different liquid viscosity methods and the compositional model was made for the 10-inch line at year 1 rates. The predicted inlet pressures and liquid viscosity for the models were:

CPTC

NOVEMBER 1994

70


Model

Inlet Pressure, psig

Liquid Viscosity at Pipeline Outlet, cp

Compositional

247

2

Black Oil, Glass Is.

301

18

Black Oil, Chew Vis.

320

29

Black Oil, Vazquez Vis.

335

62

There is a wide discrepancy between the model predictions, which can be explained by the large differences in liquid viscosity. The viscosity of the reservoir fluid was measured during the PVT analysis of the recombined wellstream. This viscosity indicated that the Glaso correlation modeled the liquid viscosity best. The poor viscosity prediction of the compositional model may be due in part to the characterization of the heavy ends of the wellstream. Errors of this magnitude in physical properties are disturbing, and merit further investigation. It was decided to use the Black Oil model with the Glaso liquid viscosity correlation for all the Pipephase runs for this example. Pipephase runs were made for the three line sizes (8, 10, and 12 inch) for the 4 production rates shown in the production forecast. The inlet pressures for the various cases were: Year No.

8 inch

10 inch

12 inch

1

530

301

199

4

433

242

164

8

194

133

119

12

126

115

110

The 8-inch line is too small to handle the rates in the early years of production. The 10inch line is at the maximum inlet pressure for the Year 1 conditions. The 12-inch line is comfortably within the maximum inlet pressure of the line. The mixture velocities at the exit of the P/P riser for the various cases are:

CPTC

NOVEMBER 1994

71


Year No.

8 inch

10 inch

12 inch

1

33.21

21.06

14.58

4

18.60

12.03

8.37

8

5.88

3.75

2.60

12

2.51

1.59

1.12

Pipephase does not print the superficial gas and liquid velocities. The values for these parameters can be calculated from the program output by the following relationships: Vsg = Ug (I-HI) Vsl = Ul HI where the Ug is the actual gas velocity, Ul is the actual liquid velocity, and HI is the slip liquid holdup. These three parameters are tabulated in the "Link" device detail report. The superficial gas velocity at the exit of the P/P riser are: Year No.

8 inch

10 inch

12 inch

1

30.01

19.00

13.14

4

15.13

9.80

6.81

8

3.90

2.47

1.71

12

1.44

0.90

0.64

The superficial liquid velocity at the exit of the P/P riser are: Year No.

8 inch

10 inch

12 inch

1

3.20

2.06

1.44

4

3.47

2.23

1.56

8

1.98

1.28

0.89

12

1.07

0.69

0.48

The value of the constant C in the erosional velocity formula can be calculated by either of

CPTC

NOVEMBER 1994

72


the following equivalent relationships: C = [Vm (ρl Vsl + pg Vsg)]0.5 = ρns0.5 Vm The values of the liquid, gas, and no-slip densities can be read from the, "Property" summary table in the Pipephase output. Substituting these values into the above equation gives the following values for C at the exit of the P/P riser: Year No.

8 inch

10 inch

12 inch

1

78.9

50.4

35.0

4

62.4

40.3

28.1

8

26.8

17.2

11.9

12

12.9

8.2

5.8

The values for C are within the API guidelines for all cases. The Taitel-Dukler-Barnea flow regimes at the end of the pipeline for the various cases are: Year No.

8 inch

10 inch

12 inch

1

Slug

Slug

Slug

4

Slug

Slug

Slug

8

Slug

Slug

Strat

12

Slug

Strat

Strat

The predicted liquid holdups (in barrels), using the Beggs and Brill correlation, are:

CPTC

Year No.

8 inch

10 inch

12 inch

1

1064

1336

1664

4

1268

1683

2197

8

1367

2018

3607

12

1563

3093

5009

NOVEMBER 1994

73


The conclusions regarding the line sizing are that the 10-inch line is the appropriate line size. The pressure drop for the 10-inch during year 1 is right at the upper limit for inlet pressure, so the production rate could be less than design if the Pipephase predictions are optimistic. The 8-inch line is definitely undersized, and the 12-inch appears oversized. The velocities in the later years of operation are a definite concern for all three-line sizes. Due to these low velocities, severe slugging is a concern during the later years. We need to quantify the occurrence and severity of slug flow. The Pipephase results indicate that the pipeline will operate in the hydrodynamic slug flow regime for all rates with the 8 inch line, years 1 to 8 for the 10 inch line, and years 1 to 4 for the 12 inch line. Year 8 for the 12-inch and year 12 for both the 10 and 12 inch lines are shown to be in stratified flow in the pipeline. This result is deceptive, since stratified flow in a downwardly sloping pipeline connected to a vertical riser is the classical severe slugging scenario. The key to estimating whether severe slugging will occur is to determine whether the Pots πss factor is less than 1. Pipephase tabulates this value as part of its "SLUG" report. The values for the various cases are: Year No.

8 inch

10 inch

12 inch

1

0.098

0.084

0.078

4

0.052

0.044

0.041

8

0.024

0.023

0.032

12

0.020

0.042

0.121

All the values of πss are much less than 1, indicating that severe slugging is possible for this pipeline. By Pots’ analysis, only those points, which are in stratified flow in the pipeline, will be in the severe slugging regime. Therefore, the predicted flow regimes for the system using the Pipephase results are:

CPTC

Year No.

8 inch

10 inch

12 inch

1

Hydrodynamic Slug

Hydrodynamic Slug

Hydrodynamic Slug

4

Hydrodynamic Slug

Hydrodynamic Slug

Hydrodynamic Slug

8

Hydrodynamic Slug

Hydrodynamic Slug

Severe Slug

12

Hydrodynamic Slug

Severe Slug

Severe Slug

NOVEMBER 1994

74


6.1.2

Slug Length Prediction

We need to estimate the slug lengths for the various cases. As mentioned in the text, the use of correlations for slug length prediction is not very accurate. Unless a transient model is available, however, these predictions are the best available method. The output for the Pipephase "SLUG" option prints tables summarizing the predictions of the Brill, et al. correlation. It also indicates an approximate slug length for severe slugging. The Brill, et al. predictions will be used as the basis for predicting slug length in hydrodynamic slug flow, and the Pots πss method will be used as the basis for predicting slug length in severe slugging. These estimates will be checked against the predictions of the Hill & Wood model for hydrodynamic slugging and the predictions of the OLGA transient model. In severe slugging, the operation is cyclic: •

The low spot fills with liquid, blocking the flow;

•

Liquid builds in the pipeline and riser until the pressure behind the liquid slug builds to a high enough level to push the slug out;

•

The liquid exits the system as a slug, increasing in velocity as the gas behind the slug expands;

•

When the gas has exited the riser, liquid will fall back down the riser, and drain in from the pipeline, starting the cycle again.

The slug length is estimated by: Ls = Lr / πss where Lr is the riser height in feet. The slug length should not vary much for a constant rate, since the operation in cyclic. The maximum slug length therefore is approximately equal to the average slug length. The values for the slug length reported in Pipephase for the 3 severe slugging conditions are:

CPTC

NOVEMBER 1994

75


a) For the 10" Line, Year 12 - Ls = 2392 ft. b) For the 12" Line, Year 8 - Ls = 3140 ft c) For the 12" Line, Year 12 - Ls = 829 ft The large decrease in the slug length for the 12" line, year 12 is due to the large increase in holdup in the line predicted by Beggs and Brill. The accuracy of this value is doubtful. For the hydrodynamic slugging cases, the "SLUG" report in Pipephase contains predictions of the average slug length and the slug length distribution using the Brill, et al. model. As detailed in the text, the Brill method is based on a very small amount of data, and it doesn’t account for most of the variables, which play a part in slug flow. To use the Brill slug length model to estimate the maximum size of a slug for equipment selection, the designer needs to estimate the slug length corresponding to a design occurrence frequency. A reasonable philosophy would be to design the separator/slug catcher for the once per year slug. Estimating the average slug frequency from the Pipephase "SLUG" report, the number of slugs per year can be calculated. The slug distribution table above can be used to estimate the slug length corresponding to one occurrence in this number of slugs per year. For example, the year 1 rate through the 10" line gives a slug frequency of 98.4 seconds by the Brill model. This frequency corresponds to (365 D/yr * 86,400 sec/D) / 98.4 sec = 320,500 slugs/yr If one slug per year can be greater than the design slug, the design slug must be at the (1 - 1/320500) * 100 = 99.9997 percentile The average slug, according to Brill, is 293 ft long. The maximum allowable slug would then be about 10.5 times as long as the average or 3080 ft long. The design percent slug was calculated for all the other hydrodynamic slug cases. The design slug was close to the 99.9999% slug length for all cases, using the one slug per

CPTC

NOVEMBER 1994

76


year criteria. Tabulating the 99.9999% data for the hydrodynamic slug cases gives design slug lengths (in feet) for the various cases of: Year No.

8 inch

10 inch

12 inch

1

2152

3153

4250

4

2079

3050

4113

8

1943

2847

3140*

12

1848

2392*

829*

* Terrain slugs To calculate the liquid surge requirements for the separator based on these slug lengths, additional information is needed. Since the slugs consist of aerated liquid, the liquid holdup in the slug must be known in order to determine the volumetric rate entering the separator. The surge volume required also depends in large measure on the capacity of the liquid dump valves on the separator. The Pipephase "SLUG" table includes a calculation of the liquid holdup in the slug, based on the Brill, et al. method. The tabulated value for the "liquid from the slug" is equal to: Volslug = Vm Hls Ap ts where Hls is the liquid holdup in the slug, Ap is the pipe cross-sectional area, and ts is the time that the liquid slug exits the pipeline. The required surge volume for the separator is: Volsurge = Volslug - ts Qdump where Qdump is the volumetric capacity of the dump valve. For this example, we will assume that the volumetric capacity of the dump valve is 1.33 times the design oil rate. This assumption gives a Qdump of 1.35 ft3/sec. The separator is also assumed to be a 2-phase separator.

CPTC

NOVEMBER 1994

77


With these assumptions, the design surge volumes (ft3) for the hydrodynamic slugging cases are: Year No.

8 inch

10 inch

12 inch

1

145

561

1398

4

195

654

1506

8

101

294

—

12

0

—

—

It is more difficult to estimate the slug volumes for the severe slugging cases using only steady state analysis. For a crude approximation of the behavior, use the following: a) Slug buildup phase When the slug is building up, there is essentially no flow out of the pipeline. The time required to build the slug is approximately: tbu = Ls/vsl The liquid rate out of the line is ~0 during this time. b) Slow gas expansion phase After the gas pressure has built to a high enough level to push the slug, the liquid rate out of the line is approximately equal to the mixture velocity, Vm. The time required for this phase is tsg. c) Fast gas expansion When the liquid slug is shorter than the riser pipe, the gas behind the slug expands rapidly. The velocities in the riser may reach 60-80 ft/sec. This time period is so much smaller than the others that it can be assumed to be virtually instantaneous. The liquid rate is so high that essentially all the liquid produced during this time goes into the separator surge. With this simplified model, the flowrates out of the pipe can be estimated. The slug cycle time (tc) is equal to tbu + tsg. The liquid produced during one cycle is equal to tc * Vsl * Ap. The liquid leaving the pipeline during the slug buildup phase is essentially 0, and the

CPTC

NOVEMBER 1994

78


volume of liquid exiting during the fast gas expansion is equal to the riser volume. The remainder of the liquid must exit during the slow gas expansion phase. Therefore, (tbu + tsg) Vsl Ap = (tsg Vm Ap) + (Lr Ap) tsg = [(tbu Vsl) - Lr] / Vsg tsg = (Ls - Lr) / Vsg To illustrate the use of these equations, we will check the outlet rates for the 10" line at year 12. The values for the key parameters are Ls = 2392 ft. Vsg = 0.90 ft/sec, Vsl = 0.69 ft/sec, Ap=0.492 ft2, and Lr = 100 ft. The value for tbu = 2392/0.69=3467 sec. The value for tsg = (2392-100)/0.90 = 2547 sec. The model indicates that the liquid flowrate out of the line is: 0 ft3/sec for 0
8 inch

10 inch

12 inch

1

145

561

1398

4

195

654

1506

8

101

294

935

12

0

49

71

Using the Brill analysis, the required surge volume for the downstream separator would be 654 ft3 for the 10" line.

CPTC

NOVEMBER 1994

79


6.1.3

Slug Frequency and Length by Hill & Wood Method

The Hill & Wood correlation for slug frequency will be used to check the slug length predictions and surge volume predictions. To use the Hill & Wood method, a calculation of the liquid holdup at the beginning of the pipeline must be made using the Taitel-Dukler stratified flow model. The method assumes a horizontal pipeline. The value for the Lockhart & Martinelli X factor can be calculated from the following equation:

V  X =  sl   Vsg 

0.9

 ρl    ρ   g

0.4

 µl    µ   g

0.1

Once X is known, the value for the liquid holdup for stratified flow can be read from Figure I:5-1. The actual gas and liquid velocities can be calculated from the superficial velocities and the stratified flow liquid holdup. Once these parameters are known, the slug frequency can be calculated. To illustrate the procedure, the slug frequency for the 10" line for the year 1 conditions will be calculated. The values for the required parameters at the beginning of the pipeline are: Vsg = 5.40 ft/sec

ρg = 1.10 lb/ft3

µg = 0.012 cp

Vsl = 2.10 ft/sec

ρl = 53.97 lb/ft3

µl = 7.1 cp

X is then: X= (2.1/5.4)°~9 * (53.97/1.1)° 4 * (7.1/.012)°~1 = 3.84 The value for the stratified flow liquid holdup, using the Taitel-Dukler model for horizontal flow, is: Hle=0.68 read from Figure I:5-1. The actual gas velocity is:

CPTC

NOVEMBER 1994

80


Ug = Vsg/(1 HIe) = 5.40/(1-.68) = 16.88 ft/sec The actual liquid velocity is: Ul = Vsl/Hle = 2.1/.68 = 3.09 ft/sec Putting these values into the Hill & Wood slug frequency equation gives: Fs = 2.74 Hle (Ug - Ul) / (D (l -HIe)) = (2.74) (0.68) (16.88-3.09)/[(9.5/12)(1-.68)] = 101 slugs/hour The length of the average slug plus bubble unit would be: Ls + Lb = (Vm) (3600)/Fs = (7.5 ft/sec) (3600 sec/hour) (1 hour)/101 slugs = 267 ft To estimate the slug length, the slug fraction must be calculated. There are several ways to do this. This example will show the easiest way to estimate the slug fraction. First, the liquid holdup in the slug will be estimated. The Gregory, Nicholson, and Aziz method is the easiest method to use, and it is as accurate as most of the methods. The liquid holdup in the slug is given by: Hls = l [l +(V,,’/28.4)1.39] = 1/[1+(7.5/28.4)1.39] = 0.864 The liquid holdup in the bubble can be assumed to be ~0.20 for much of the slug flow range. From a material balance, the slug fraction can be calculated from these values and the overall liquid holdup prediction: SF = (Hl - Hlb) /(Hls - Hlb)

CPTC

NOVEMBER 1994

81


Hl = 0.42 by the Beggs and Brill method at the beginning of the pipeline. Therefore, SF = (0.42 - 0.20)/(0.864 - 0.20) = 0.33 Since the slug fraction is equal to the slug length divided by the slug length plus bubble length, the slug length is equal to: Ls = SF (Ls + Lb) = 0.33 x 267 = 88ft. This value for the average slug length is much less than the 293 ft average slug length calculated by the Brill et al. method. The average slug lengths for the 10" line are: Year Number

Average Slug Length, ft.

1

88

4

52

8

49

12

Terrain Slugging

The Hill and Wood method does not predict a slug length distribution. Experimental evidence from the Prudhoe Bay field tests and BHRG laboratory studies indicate that the maximum slug length is approximately 6 times the average slug length. Using this value, the design slug lengths are:

CPTC

NOVEMBER 1994

82


Year Number


1

528

4

312

8

294

12

2392

The design surge volume would be: Volsurge = (Lsdes Hls Ap) - [(Lsdes Qdump) / Vm] For the year 1 conditions, this volume is: Volsurge = [(528 ft) x (.864) x (0.492 ft2 )] - [(528 ft) x (1.35 ft3/sec) / (7.50 ft/sec)] = 129ft3 The values of the design surge volume for the 10" line during all the years are: Year Number


1

129

4

76

8

2

12

49

The design surge volume using the Hill & Wood method of 129 ft3 is considerably less than the 654 ft3 calculated by use of the Brill et al. method. For comparison, the liquid volume needed for separation is about 200 ft3, based on a 3 minute liquid residence time in the vessel. The transient OLGA program was run for the four design cases for the 10" line. The inlet pressures using OLGA for the cases were:

CPTC

NOVEMBER 1994

83


Year Number

OLGA

Beggs and Brill

1

228 psig

301 psig

4

187 psig

242 psig

8

115 psig

133 psig

12

101-113 psig

115 psig

The pressure drops calculated by OLGA are much less than the Beggs and Brill pressure drops. This is due in part to the physical properties used in the OLGA runs, since OLGA uses a compositional model similar to Pipephase’s compositional model. Most of the difference, however, is due to the different multiphase flow models. OLGA tends to predict pressure drops that are too low in slug flow, and Beggs and Brill tends to overpredict pressure drops. The actual inlet pressure is between the two answers. For year 1, for instance, the expected inlet pressure would probably be about 270-280 psig. The flow regimes predicted by OLGA were: Year Number

Flow Regime

1

Hydrodynamic Slug Flow

4


8


12

Terrain Slug Flow

OLGA confirms the Pipephase flow regime predictions, which indicated the line should be in slug flow for years 1 through 8, and then experience terrain slugging in year 12. The liquid holdup predictions of OLGA were:

CPTC

Year Number

OLGA

Beggs and Brill

1

1365 bbl

1336 bbl

4

1831 bbl

1683 bbl

8

2565 bbl

2018 bbl

12

2915 bbl

3093 bbl

NOVEMBER 1994

84


The liquid holdup predictions of OLGA and Beggs and Brill are quite close. Beggs and Brill’s holdup predictions in slug flow are pretty good, as seen above, but they can be very poor for stratified flow, as shown in Example 2. For the year 12 conditions, OLGA shows a cyclic terrain slugging pattern, as seen in Figures 1:6-1 to I:6-4. The results shown for about the first 15,000 seconds reflect the initial conditions selected for the problem, and can be ignored. After 15,000 seconds, a cyclic pattern of terrain slugs is seen. Every 2400 seconds, a slug exits the pipeline with a volume of about 11 cubic meters (-70 bbls). This slug volume corresponds to a slug length of ~790 feet, which is considerably smaller than the 2392 feet predicted by the Pots method. The OLGA results are much more believable than the Pots predictions. The pressure oscillates between 101 psig and 113 psig. The liquid mass flow rate out of the pipeline peaks at a rate of about 40 kg/sec, which corresponds to an instantaneous flowrate of 21,700 BPSD, compared to the average flowrate of 5203 BPSD. The gas flow rate peaks at about 0.28 kg/sec, which is equivalent to 1.0 mmSCFD, compared to its average rate of 0.26 mmSCFD. The slug-tracking model in OLGA V3.1 was run for the three years that were in hydrodynamic slug flow. The slug-tracking model tracks each individual slug, and determines whether it grows, dissipates, or merges with another slug as it flows down the line. The slug-tracking runs for these cases, however, were unable to converge on a value for the average slug length. For year 1, for example, runs were made with three different numbers of pipeline segments. If the model converges on a solution, the slug length would be independent of the number of segments used. For this case, the average slug length vs. number of pipeline segments was:

CPTC

Number of Segments

Average Slug Length, feet

79

197

154

82

294

51

NOVEMBER 1994

85


The maximum number of pipeline segments in the commercial version of OLGA is 297. Since the slug length was still decreasing using the maximum number of segments, the program had not reached a value for the slug length that was independent of pipeline discretization. As a result, the only conclusion that can be drawn from the slug-tracking analysis is that the average slug length is probably less than 50 feet for the year 1 conditions. The slug-tracking predictions are in line with the results of the Hill and Wood model. Both predictions are much less than the Brill, et al. method. In summary, the OLGA runs verified that the 10" pipeline should be capable of flowing all rates within the allowable pressure drop. They confirmed that the pipeline would be in hydrodynamic slug flow for the first 8 years, and then operate in the terrain slug flow regime by year 12. The slug lengths predicted by Pipephase for the terrain slugging case were found to be too large, and the slug lengths for the hydrodynamic slugging cases were predicted to be of the same magnitude as the Hill and Wood predictions. If this example was an actual design, checks of other transient phenomena such as startup, shutdown, rate changes, etc. should be performed to determine whether they would cause slugs which would govern the size of the slug catcher. Based on this analysis, the slug catcher should be designed for a surge volume of about 150 ft3, based on the Hill and Wood results. 6.2

EXAMPLE 2 - Gas/Condensate Line

A pipeline is needed to flow the production from a subsea well development to shore. The offshore field produces a mixture of gas, water, and condensate. Glycol is injected into the line for hydrate suppression. The water quantity is about 4 bbl/mmSCF, and the condensate production rate is also about 4 bbl/mmSCF. The peak production rate is 500 mmSCFD of gas and condensate. The minimum turndown is 25% of design rate. The arrival pressure onshore is 900 psia. The inlet pressure to the line cannot exceed 1200 psia under normal operation. The wellstream composition, including water and injected glycol is given in Table 1.

CPTC

NOVEMBER 1994

86


The pipeline elevation profile is shown in Table II. The wellhead pressure will decrease with time as the reservoir becomes depleted. Initially, the wellhead pressure will be over 4000 psig. For design, the wellhead temperature and pressure are estimated to be 222 degrees F and 1215 psia, simulating conditions late in the life of the field. The pipeline is unburied, and is coated with 3 inches of concrete and 0.125 inches of epoxy coating. 6.2.1 TABLE I, Wellstream Composition Component

Mole %

N2

3.68468

CO2

11.11650

C1

75.79904

C2

3.82831

C3

1.15655

iC4

0.16298

nC4

0.23880

iC5

0.08426

nC5

0.06506

C6’s

0.05431

C7’s

0.07390

C8’s

0.10359

C9’s

0.07833

C10’s

0.03915

C1l’s

0.02022

Cl2’s

0.01223

CPTC

NOVEMBER 1994

87


Cl3’s

0.00985

C14's

0.00794

C15's

0.00556

C16’s

0.00460

Cl7’s

0.00248

C18's

0.00157

C19's

0.00126

C20's

0.00074

C21's

0.00073

C22's

0.00054

C23's

0.00037

C24's

0.00031

C25's

0.00024

C26's

0.00019

C27’s

0.00016

C28+

0.00021

Ethylene Glycol

0.12530

Water

3.32004 Totals

6.2.2

100.00000

TABLE 11, Pipeline Elevation Profile

Cumulative Distance, km 0.0 0.35

Water Depth, m 216 218

Segment Length, ft

Elevation Change, ft

1148

-6.6

1.45 1.75 2.25

218 219 219

3608 984 1640

0.0 -3.3 0.0

CPTC

NOVEMBER 1994

88


2.55 3.15 3.45 3.95 4.30 6.35 6.90 7.30 7.75 7.95 8.30 8.90 9.70 10.00 10.40 11.20 12.40 13.45 14.45 15.70 16.95 18.45 20.20 22.10 24.10 25.85 30.25 34.60 42.10 49.55 56.15 60.45 61.75 63.05

CPTC

218 218 219 219 218 218 217 217 218 218 216 214 210 208 208 200 190 180 170 160 150 140 130 120 110 100 90 80 70 65 65 65 60 50

984 1968 984 1640 1148 6724 1804 1312 1476 656 1148 1968 2624 984 1312 2624 3936 3444 3280 4100 4100 4920 5740 6232 6560 5740 14432 14268 24600 24436 21648 14104 4264 4264

NOVEMBER 1994

3.3 0.0 -3.3 0.0 3.3 0.0 3.3 0.0 -3.3 0.0 6.6 6.6 13.1 6.6 0.0 26.2 32.8 32.8 32.8 32.8 32.8 32.8 32.8 32.8 32.8 32.8 32.8 32.8 32.8 16.4 0.0 0.0 16.4 32.8

89


74.25 78.50 79.45 80.85 81.95 82.60 83.40 84.10

6.2.3

40 30 25 20 15 10 5 0

36736 13940 3116 4592 3608 2132 2624 2296

32.8 32.8 16.4 16.4 16.4 16.4 16.4 16.4

Pipephase Simulation Comparisons

Pipephase simulations of the operation were made for 3 rates (125, 250, and 500 mmSCFD) for two line sizes (30" and 28"). Since this line is a gas-condensate line, the Oliemans method was used to estimate the pressure drop in the system. The TaitelDukler-Barnea method was used to predict the flow regime. As a check, the Beggs and Brill method was run for the selected line size. A transient analysis of this line was also performed using OLGA. It will therefore be possible to compare the predictions of Oliemans and Beggs and Brill against OLGA. The inside diameters used for the 30" line was 28.5", and the inside diameter for the 28" line was 26.5". The predicted inlet pressures using the Oliemans correlation were: Production Rate, mmSCFD

30” Line

28” Line

125

932 psig

938 psig

250

970 psig

994 psig

500

1120 psig

1200 psig

CPTC

NOVEMBER 1994

90


The 30" line is comfortably within the wellhead pressure specification. The 28" line is right on the specification at design rate. A check of the pressure drop for the 28" line using the Beggs and Brill correlation gives: Production Rate, mmSCFD

Inlet Pressure, psig

125

947

250

1008

500

1226

Beggs and Brill indicates a higher pressure drop than Oliemans, making the 28" line more questionable. Both line sizes will continue to be checked. The OLGA transient simulations will be used as a third check of the pressure drop to determine whether the 28" line appears acceptable. The total liquid holdup for the line (barrels) using Oliemans and Beggs and Brill were: Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

125

11,760

9,480

14,560

250

7,890

6,280

13,100

500

5,000

4,230

12,340

mmSCFD

The superficial gas and liquid velocities are not printed in the Pipephase output. The superficial velocities can be calculated from the actual gas or liquid velocities and the liquid holdup. For gas/condensate lines, this is difficult because of the poor formatting of the liquid holdup in the output table. Pipephase prints the value of the liquid holdup as a two decimal place value (e.g. 0.02), which is not adequate to accurately determine the superficial velocities. For holdup below 0.005, Pipephase prints 0.00. Fortunately, it is possible to calculate accurate values for the superficial gas and liquid velocities and liquid

CPTC

NOVEMBER 1994

91


holdup from the information give in the output table. The superficial liquid velocity can be calculated from the formula: Vsl = [Ul(Ug - Vm)]/(Ug- Ul) where Ug = Actual gas velocity, ft/sec Ul = Actual liquid velocity, ft/sec Vm = Mixture velocity, ft/sec The superficial gas velocity can be calculated by: Vsg = Vm - Vsl The liquid holdup is given by: HI = Vsl / Ul The superficial gas velocities (ft/sec) for the various cases are: Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

mmSCFD 125

4.66

5.39

5.38

250

9.32

10.77

10.77

500

18.54

21.40

21.38

The superficial liquid velocities (ft/sec) are: Production Rate,

Oliemans - 30”

Oliemans - 28”

0.0189

0.0224

Beggs-Brill - 28”

mmSCFD 125

CPTC

NOVEMBER 1994

0.0226

92


250

0.0381

0.0439

0.0471

500

0.0787

0.0884

0.0895

The values of C in the erosion velocity for the various cases are:

Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

125

0.0189

0.0224

0.0226

250

0.0381

0.0439

0.0471

500

0.0787

0.0884

0.0895

mmSCFD

All the C values are well below 100, so corrosion-erosion should not be a concern. The dominant flow regimes for the various cases, using the Taitel-Dukler-Barnea method are: Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

mmSCFD 125

Slug

Slug

Slug

250

Strat./Slug

Strat./Slug

Strat./Slug

500

Strat./Annul.

Strat./Annul.

Strat./Annul.

The line is shown to be in slug flow at the 125 mmSCFD rate. Some of the segments of the line are shown to be in slug flow at 250 mmSCFD, while others are shown to be in stratified flow. For the 500 mmSCFD rate, all the segments are shown to be in either stratified or annular flow. Pipephase will not print a slug report for the runs in which the Oliemans correlation was used. It did print slug lengths for the Beggs and Brill runs. The slug lengths predicted by

CPTC

NOVEMBER 1994

93


the Brill et al. method are 1460 to 1580 ft average slug length for the 28" line. These numbers are very unrealistic for gas/condensate systems of this diameter. The design criterion for the slug catcher surge volume is usually the liquid volume associated with pigging the pipeline. This volume can be estimated from the Pipephase output. The pig travels at approximately the mixture velocity, Vm. The average values for Vm (ft/sec) for the various cases is: Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

mmSCFD 125

4.87

5.62

5.59

250

9.54

10.92

10.79

500

17.61

19.54

19.28

The pipeline is 84.1 km or 275,850 ft. long. The pig traverse times (seconds), therefore, are: Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

125

56,640

49,080

49,350

250

28,920

25,260

25,570

500

15,660

14,120

14,310

mmSCFD

An approximation of the volume of the slug ahead of the pig is that the slug volume is equal to the liquid holdup minus the amount of liquid produced during the pig traverse time. Numerically, this equates to: Volslug = (Hl Ap Lp) -( tpig Vsl Ap)

CPTC

NOVEMBER 1994

94


The slug volumes (barrels) for the various cases are: Production Rate,

Oliemans - 30”

Oliemans - 28”

Beggs-Brill - 28”

125

10,920

8,730

13,800

250

7,020

5,520

12,280

500

4,030

3,380

11,470

mmSCFD

If transient analysis is not available, the table above would be used to design the surge volume required for the slug catcher. For the 28" line, the slug catcher would be sized for the 13,800 bbl surge volume, with the recognition that neither Oliemans nor Beggs-Brill predicts liquid holdup well for gas/condensate pipelines. We will use the OLGA transient program to check the design of the line to confirm the line size, determine whether slugging is likely, and size the slug catcher. OLGA simulations were run at the three design rates for both the 28 inch and 30 inch lines. The predicted inlet pressures at the various rates were: Rate, mmSCFD

30 Inch

28 Inch

125

1045 psig

1018 psig

250

980 psig

982 psig

500

1100 psig

1179 psig

A comparison of the OLGA, Oliemans, and Beggs and Brill predictions for the 28 inch line at the various rates is:

CPTC

NOVEMBER 1994

95


Rate, mmSCFD

OLGA

Oliemans

Beggs-Brill

125

1018 psig

938 psig

947 psig

250

982 psig

970 psig

1008 psig

500

1179 psig

1200 psig

1226 psig

The Oliemans predictions match the OLGA predictions quite closely, except at the lowest rate. The difference between Oliemans and OLGA at the 125 mmSCFD rate is due to the high liquid holdup predicted by OLGA, as shown below. The OLGA results confirm that the pressure drop in the pipeline is less than the allowable drop for the 28" line, so the 28" line can be used. The predicted liquid holdups from OLGA are: Rate, mmSCFD

30 Inch

28 Inch

125

52,800 bbls

32,900 bbls

250

8,800 bbls

3,790 bbls

500

3,310 bbls

2,650 bbls

Comparing these results with Oliemans and Beggs and Brill for the 28 inch line shows the wide disparity in results between the methods: Rate, mmSCFD

OLGA

Oliemans

Beggs-Brill

125

32,900 bbls

9,480 bbls

14,560 bbls

250

3,790 bbls

6,280 bbls

13,100 bbls

500

2,650 bbls

4,230 bbls

12,340 bbls

These results are plotted in Figure I:6-5. OLGA’s holdup predictions are much higher than either Oliemans or Beggs and Brill at the lowest flowrate, and they are much lower at the other rates. The OLGA holdup jumps when the gas velocity is too low to pull the liquid up the upwardly inclined pipeline segments. OLGA indicates that, for these operating

CPTC

NOVEMBER 1994

96


conditions, this minimum superficial gas velocity is about 9 ft/sec, corresponding to a throughput rate of about 220 mmSCFD. Below this rate, the liquid holdup builds up rapidly. This behavior has been observed in the field and in experimental studies in inclined flow. Neither Beggs and Brill nor Oliemans model this sensitivity to flowrate in inclined pipes. The predicted flow regimes from the OLGA simulations were: Rate, mmSCFD

30 Inch

28 Inch

125

Stratified and Slug

Stratified and Slug

250

Stratified

Stratified

500

Stratified

Stratified

OLGA shows that some of the pipeline segments are in slug flow at the 125 mmSCFD rate for both line sizes. For the higher rates, all of the segments are shown to be in stratified flow. The comparison between the OLGA results and the Taitel-Dukler-Barnea predictions is: Rate, mmSCFD

OLGA

Taitel et al.

125

Strat./Slug

Slug

250

Strat.

Strat./Slug

500

Strat.

Strat./Annul.

Although the comparison might not seem very good, the results of the two methods are fairly close. OLGA predicts a somewhat lower transition velocity from Stratified to Slug flow than Taitel-Dukler-Barnea, so that it shows slug flow only at the lowest rate. At the highest rate, the predictions are really the same, since OLGA shows stratified flow with considerable entrainment. Visually, this is indistinguishable from annular mist flow. The only difference between the two methods is at the 250 mmSCFD. A good rule of thumb is that slug flow does not occur for liquid holdups less than about 20%. The liquid holdup by both the Oliemans and Beggs and Brill methods for the 250 mmSCFD rate is <7%, so the

CPTC

NOVEMBER 1994

97


prediction of slug flow by Taitel, et al. doesn’t appear likely. The OLGA prediction, therefore, is more believable for this rate. OLGA simulations for pigging were also performed. The liquid slug sizes ahead of the pig for the various rates in the 28" line are: Rate, mmSCFD

Liquid Slug Size, bbls

125

33,900

250

3,770

500

1,920

If the line is run to equilibrium at the lowest rate, the slug catcher would have to be sized for about 40,000 bbls. The capital expenditure to build a slug catcher that large would be very significant. The best solution for the slug catcher design probably is to size the slug catcher for the 250 mmSCFD rate, and pig the line frequently when the rate is less than 250 mmSCFD. To get an estimate of the pigging frequency that would be required, calculate the time required to produce ~3770 bbls of liquid from the wellstream at the desired rate. At 125 mmSCFD, the pigging frequency would be: 3770 bbls * 42 gal/bbl / 7.48 gal/ft / Vsl ft/sec / Ap ft2 / 3600 sec/hr / 24hr/D = 3770 * 42 / 7.48 / 0.0224 / 3.83 / 3600/ 24 = 2.86 days If the line is pigged every 2-3 days at the 125 mmSCFD rate, the smaller slug catcher could be used. The design volume for the slug catcher should comprehend errors in the composition, pipeline topography, and errors in the multiphase flow methods. The amount of the increase should be based on sound engineering judgment. For our example, the 3770 bbl maximum slug size calculated for the 250 mmSCFD rate should be increased to a design slug catcher surge volume of at least 5000 bbls.

CPTC

NOVEMBER 1994

98


In summation, the best design for this pipeline would be: a) install a 28" line; b) install a slug catcher with a 5000 bbl surge capacity; c) pig the line every 2 days at production rates less than 250 mmSCFD.

CPTC

NOVEMBER 1994

99


Figure I: 6-1 Liquid Holdup for Example 1, Year 12

CPTC

NOVEMBER 1994

100


Figure I: 6-2 Inlet Pressure for Example 1, Year 12

CPTC

NOVEMBER 1994

101


Figure I: 6-3 Liquid Flowrate Out of line, Example 1, Year 12

CPTC

NOVEMBER 1994

102


Figure I: 6-4 Gas Flowrate Out of line, Example 1, Year 12

CPTC

NOVEMBER 1994

103


Figure I: 6-5 Liquid Holdup Predictions for Example 2

CPTC

NOVEMBER 1994

104


SECTION 7.0 - REFERENCES Barnea, D. “A Unified Model for Predicting Flow-Pattern Transitions for the Whole Range of Pipe Inclinations,” Int. J. Multiphase Flow 13, No. 1, 1-12,, 1987. Bendiksen, et al., “The Dynamic 2-Fluid Model OLGA: Theory and Application,” SPE Production Engineering, pg. 171, May 1991. Brill, J.P., et al., “Analysis of Two-Phase Tests in Large-Diameter Flow Lines in Prudhoe Bay Field,” Society Petroleum Engineering Journal, pgs. 363-347, June 1981. Hill, T.J. and Wood, D.G., “A New Approach to the Prediction of Slug Frequency,” Society of Petroleum Engineers, pgs. 141-149, SPE 20629, 1990. Taitel, Y. and Dukler, A.E., “A Model for Predicting Flow Regime Transitions in Horizontal and Near Horizontal Gas-Liquid Flow,” AIChE Journal, 22, No. 1, pgs. 47-55, 1976.

CPTC

NOVEMBER 1994

105

deepstar_chevron-multiphase-design-guide-pt-1

Recommend Documents