Feature Report Engineering Practice
Reduce Gas Entrainment In Liquid Lines P0
T 0
V-1
h
P1, T 1
Follow these tips to properly size self-venting lines and vortex breakers
H 0
P2, T 2
H 1
V-2
H 2
Tamagna T amagna Ukil and Thomas Mathew Reliance Industries Ltd.
A
Ground
common practice in the chemical process industries (CPI) is to implement so-called valuemaximization projects (VMP) to increase production or reduce production costs in order to increase profit margins. With such projects, one main objective of the design team is to incur minimum capital expenditures. Because most VMPs aim to increase throughput or production yield, many such projects involve changes to the process that result in an increase in the volume of feed flowing into a gasliquid separator (GLS). The system modifications that are required often call for: t5IFEFTJHOPGBOFX(-4UPBDDPN5IFEFTJHOPGBOFX(-4UPBDDPNmodate the increased flow, or t5IFNPEJGJDBUJPOPGUIFWFTTFMJOUFS5IFNPEJGJDBUJPOPGUIFWFTTFMJOUFSnals and associated piping to handle the increased feed flow Increased feed flow into any GLS can lead to the entrainment of gases into the liquid lines. Such gas entrainment can lead to pulsating flows in the line, which can result in vibration and potentially destabilize the downstream processes. In many cases where GLS are provided with “gravity-flow pipelines” — a common approach, as it pro vides an inexpensive inexpensive way to transp transport ort liquids — the use of self-venting pipelines coupled with properly sized vortex breakers can mitigate the problem of entrainment of gases into liquid lines.
(V-1) is P0 (psig) and its operating temperature is T 0 (°F). The operating pressure and temperature of the second vessel (V-2) are P 2 and T 2, respectively. The pressure and temperature of the liquid at the exit nozzle of V-1 are P are P1 and T 1, respectively. In Figure 1, the region from the exit of V-1 to the inlet of V-2 is highlighted with a dashed outline. It shows that the associated piping of the system consists of pipes and elbows. The following assumptions are considered for this system: t -JRVJEGMPXJOHUISPVHIUIFMJOFJT incompressible t 5IFT 5IFTZTUFN ZTUFNJTJO JTJOTUFBE TUFBEZTUB ZTUBUF UF t 5IFSF 5IFSFJTOP JTOPGMBT GMBTIJOHP IJOHPGMJR GMJRVJE VJE t 1SFTTVSFT 1SFTTVSFT P P0, P1 and P 2 are constant t 5IF 5IFQJ QJQF QFTJ[ TJ[FJ FJTV TVOJG OJGPSN PSN
System equations
FIGURE 1. Shown here is a typical gas-liquid separator, with gravity fow rom V-1 to V-2 [ 1] e p i p f o t o o f r e p r a e y r e p r a l l o D
20
15
10
5
0
1
2 3 4 6 Nominal pipe size, t
8
FIGURE 2. The relationship between amortized capital cost per oot o pipe and nominal pipe size is shown here [3]
approach as stated by Moharir [ 3 3], ], the cost of the pipe material per unit length for a run of pipe with diameter D is calculated using Equation (1):
Step 1. The pipeline is sized for liquid flow using a conventional line-sizing (1) approach for typical velocity considerations and least annual cost. Table 1 Along with the pipe, the cost of access accessooshows typical liquid velocities in steel ries and fittings must also be factored pipelines. in, hence their number must also be Table 1 shows typical velocities in computed on a per-unit-length basis. steel pipelines with liquid flow [ 3]. 3]. For instance, if a pipeline of 100 ft has It provides a good estimate for the 5 gate valves, 4 long-radius elbows of preliminary selection of the pipeline 90 deg, 2 tees and 7 weld joints, then its size with respect to its nominal bore per-unit fitting cost can be taken col(N.B.) dimensions. As Table 1 provides lectively as a factor F factor F . If the amortizageneralized data, readers can use the tion rate is A is A M and the annual mainteTheoretical basis values provided for any type of pipes, nance cost is a fraction G of the capital A typical GLS arrangement with irrespective of metallurgy or material cost, then the annualized capital plus maintenance cost of the pipeline, C P, is gravity flow is shown in Figure 1. The of construction. is By applying the lowest-annual-cost calculated using Equation (2): operating pressure of the first vessel 42
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TABLE 1. TYPICAL VELOCITIES IN STEEL PIPELINES WITH LIQUID FLOW [3 ]
Start
Nominal pipe size, in.
2 or less
3 to 10
10 to 20
Liquid and line
Velocity, ft/s
Velocity, ft/s
Velocity, ft/s
Pump suction
1 to 2
2 to 4
3 to 6
Pump discharge (long)
2 to 3
3 to 5
4 to 7
Discharge heads (short) 4 to 9
5 to 12
8 to 14
Boiler feed
5 to 9
5 to 12
8 to 14
Drains
3 to 4
3 to 5
Water Obtain initial pipe diameter using Table 1
Optimize the diameter using annual cost approach to get D
'
NO '
Check D < 0.31
Select the N.B. o pipe such that D < 0.31 '
YES Select the D and size vortex breakers '
Sloped sewer
——
3 to 5
4 to 7
(Normal viscosities)
1.5 to 2.2
2 to 4
3 to 6
Pump suction
2.5 to 3.5
3 to 5
4 to 7
Discharge heads (long)
4 to 9
5 to 12
8 to 15
Boiler feed
3 to 4
3 to 5
Hydrocarbon liquids
Drains
Vortex breakers to be o 2D X 2D dimension
FIGURE 3. This fowsheet illustrates the types o decisions that must be made to properly size gravity fow lines and vortex breakers, to reduce gas entrainment
(2) Rearranging Equations (1) and (2) produces Equation (3):
(3) In most cases, another component, C F , is needed to calculate is the operating cost. However, in this case, the operating cost component C F is not considered due to the absence of any rotary equipment. Differentiating C P with respect to D, to obtain optimum diameter of the pipeline ( D ) and setting it to zero, Equation (3) can then be simplified as follows:
——
——
——
Pump suction
——
——
——
Medium viscosity
——
1.5 to 3
2.5 to 5
Tar and fuel oils
——
0.4 to 0.75
0.5 to 1
Discharge (short)
——
3 to 5
4 to 6
Drains
1
the fluid inside a vessel does not rotate and if the liquid level in the vessel is below a certain height, then gas will get sucked into the liquid line. A conservative estimate of this level was derived by Harleman et al. [1], Harleman’s equation is:
(5)
Equation (5) can be used to estimate the height of the liquid inside V-1 below which the gas would be sucked into the liquid line. Experiments on 13/16-in. pipeline and on 1-in. to 4-in. pipelines by Simpson and Webb [ 2], respectively, show (4) that if the Froude number in the pipeFigure 2 shows the relationship be- line is less than 0.31, then gas will not tween the amortized annual cost per be entrained. If the Froude number unit length of pipe (ft) and nominal of the liquid flowing in the pipeline pipe size (nominal bore). is greater than 0.31, then gas starts From the two methods described getting swept up by the liquid. High, above, D is obtained as an initial two-phase pulsating flow is observed line size in terms of nominal pipe size when the Froude number is between (nominal bore) of the pipe. 0.31 and 1. Step 2. The next step is to carry out This is the basis of design for selfthe Froude number analysis for the venting lines: Any provision for selfline using the diameter obtained from venting lines should ensure that the Step 1. As per Simpson’s article [ 2], if Froude number remains between 0 b
——
Viscous oils
Stop
b
——
1.5 to 3
——
and 0.31. The typical velocity of liquid in self-venting pipelines is in the range of 1 ft/s. Step 3. When the flow inside a vessel is rotational, vortex breakers should be provided to prevent gas entrainment into liquid lines. If V-1 has a feed entry point that is tangential to the vessel, it will induce a swirling motion in the liquid, like a whirlpool. If this swirling motion is strong enough to reach the liquid exit nozzle of V-1, then it would lead to entrainment of gas into the liquid pipeline. Borghei’s experiments [ 4] in pipelines of 2-in. to 4-in. show that vortex breakers with dimensions double the nominal bore of the pipe are highly efficient in reducing the vortex effect inside the vessel. Thus in V-1, with a self-venting liquid exit line, the vortex breaker arrangement should be in the form of a cross (+). When the vertical and horizontal dimension of the plates that are used to fabricate the vortex breaker have a dimension of 2 D’ , each can substantially reduce the entrainment of gas into the liquid exit. The steps described above can be summarized in the flowsheet shown in Figure 3.
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Engineering Practice Authors
NOMENCLATURE
A M C D C P F G g h
Amortized cost per unit length of Initial pipe dia., in. D pipe, $/ft D’ Optimum pipe dia., in. Cost per unit length of pipe, $/ft D” Dia. of self-venting line, in. Total capital cost per unit length of P 0 , P 1, P 2 Pressure shown in Fig. 1, psig pipe, $/ft T 0 , T 1, T 2 Temperature shown in Pipe fitting cost per unit length of Figure 1, °F pipe, $/ft Velocity of the liquid through the V Maintenance cost per unit length of pipeline, ft/s pipe, $/ft Cost per unit length of 2-in. nomi X Acceleration due to gravity, ft/s2 nal bore pipe of the same material Height inside V-1, ft and schedule, $/ft
The following conclusions can be made from the discussion above: 1. The line size full of liquid will always be smaller than the self-venting line. 2. The work described in Refs. 2 and 4 are based on small lines (up to 4-in. nominal bore). 3. If liquid flow varies during operation, the pipe should be sized to accommodate the maximum possible flow.
bb
4. D obtained from Equation 5 should be rounded off to the higher nominal bore of pipe of standard available size. ■ Edited by Suzanne Shelley
References 1. Yu, F.C., Hydrocarbon Proc., Nov. 1997. 2. Simpson, L.L., Chem. Eng., June 17, 1960, p. 191. 3. Moharir, A.S., Pipe hydraulics and sizing, IIT Bombay, May 7, 2008. 4. Borghei, S.M. Partial reduction of vortex in vertical intake pipe, Scientiairanica, Vol 17, Issue 2.
Tamagna Ukil is the Manager of PTA-Process at Reliance Industries Ltd. (Reliance Corporate Park, Ghansoli, 7-B Ground Floor, Navi Mumbai Maharashtra, India; Phone: +912-244-783-452; Email:
[email protected]). He holds a B.S.Ch.E. from Utkal University. He is a Certified Piping Engineer from IIT Bombay, and has been working with Reliance Technology Group, PTA Division, to provide advanced technical services in the field of design, simulation and process optimization for the manufacture of purified terephthalic acid (PTA). Thomas Mathew is president of Reliance Industries Ltd. He graduated as a Chemical Engineer from Kerala University (Trichur Engineering College), and spent the first 16 years of his career involved in the production of ammonia from numerous raw materials, including natural gas, naphtha, fuel oil and coal. Mathew participated in the startup of two coal gasification plants and served as plant manager for five years in the coal gasification plant at Ramagundam, India. He joined Reliance in 1985 and took charge of the commissioning and startup of several petrochemical plants, before heading the manufacturing operations of the Reliance’s Patalganga Complex. He leads the Centre of Excellence in PTA and Gasification within Reliance.
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