Flow Rate Vs Pressure Loss 4

[asade]

Good day All,

Please, I need insight on how to achieve a low pressure at a pipeline tie-in point. The case at hand is stated below:
A crude oil pipeline pump is to transport a 70000 bopd from an offshore storage facility through a departing pipeline and tie-in to another pipeline which has been derated to 285 psig. Based on the pump performance curve, the discharge pressure from the oil pipeline pump is 526 psig which is greater than the MAWP of the tie-in pipeline.
The oil flow rate is to be maintained while the delivery pressure at the tie-point should not exceed the MAWP of the pipeline which is 285 psig. Is it possible to maintain the flow rate at low pressure at the tie-point?

I developed a 2 case models using PIPEPHASE software. The cases are to see the effect of installing a PCV to regulate the downstream pressure and without a PCV.

Case 1: Regulating the downstream P by setting the PCV @200 psig gives a delivery pressure of 173.6 at the tie-point. Discharge pressure from the pump is 553.2 psig with pump power of 787.6 psig and 0.85 efficiency.

Case 2: The delivery pressure at the tie-in point is 526.3 psig without any PCV installed downstream on the oil pipeline pump discharge line.

Please, I need your comments on the two results I provided.

Also, a colleague suggested that we should install a PSV on the departing pipeline to provide pressure relief should the pressure goes above the required downstream tie-in pressure. The departing pipeline MAWP can withstand any excess pressure that could be generated from the pump. Is it logical to install a PCV on the line with the aim to control the tie-in pont pressure which is some kilometers away?

Thanks for your anticipated comments.

 

[SNORGY]

The system losses are controlling the pumping pressure, not the pump. Based on the information you provided, you appear to need a substantial amount of throttling to achieve the tie in pressure, and without this throttling, you arrive at the tie in point after a pressure loss of 26.9 psi. It looks like in either case your pipeline losses are on the order of 26.4 to 26.9 psi. If that is true then you do not need the 526 psig pump. The pump appears to be overkill for the application. Choose a lower head, lower power pump.

I assume by a pumping power stated as 787.6 psig you meant to say 787.6 hp.

With respect to the second question, it is common to have pressure control and over-pressure protection in a pipeline, each separate from the other. In fact, in Canada, it is a pipeline code requirement, although the over-pressure protection does not necessarily need to take the form of a relief valve. I am not familiar with B31.4 & B31.8 but I would suspect the requirements are similar.

 

[zdas04]

Over pressure in this case is one of the (many) situations where you just do the credible scenario analysis and let that guide you. For example, if the only credible source of over pressure is the pump, then you can put a high-pressure kill for the pump on the pipeline downstream of the pressure regulating valve and be done. If there are other credible scenarios, then a PSV may be needed. Or not. Just do the analysis for that system and make a rational judgement.

There are times that it makes sense to leave an over-powered pump in service and dump excess pressure across a pressure regulator. For example, if a replacement pump has too long a lead time or the pressure requirement is scheduled to change later, etc. All things considered it may make economic sense to follow SNORGY's recommendation and get the right pump. It may not.

Rules of thumb are a crappy substitute for Engineering analysis in either case.

David Simpson, PE
MuleShoe Engineering

"Belief" is the acceptance of an hypotheses in the absence of data.
"Prejudice" is having an opinion not supported by the preponderance of the data.
"Knowledge" is only found through the accumulation and analysis of data.

 

[asade]

Thanks SNORGY and zdas04 for the insight.

On the pump power, I meant 787.6 hp not "787.6 psig".

The oil pipeline pump is an existing pump. The challenges occurred because the tie-in pipeline was derated to lower pressure when the production facility was shut-down.

 

[BigInch]

Yes you can and you will need to cut the discharge pressure on the pump discharge. Assuming that the tie-in is offshore and that a PCV at the tie-in itself is not a practical solution, the PCV will have to be at the pump discharge itself.

High pressure kill switches and other forms of "electronic pressure control", including the so-called High Integrity Protection Systems, "HIPS", while often found in chemical process plants, are not considered a valid mode of pressure control in pipeline work. Pipeline codes generally do not recognize kill switches as a form of overpressure protection (in case the discharge PCV fails), so a hard-piped PSV is needed. Many companies have HIPS prohibitions specifically written into their pipeline design standards as well, Saudi Aramco being one I can name in particular.

If it ain't broke, don't fix it. If it's not safe ... make it that way.

 

[asade]

BigInch,

The PCV is installed at the pump discharge. The pump is located on an offshore platform.

Yes, I agreed with you on the fact that the PCV can fail though another PCV bypass connection was designed. It was the case of PCV failure that makes us consider the installation of PSV on the departing pipeline from the platform as an option to prevent the downstream tie-in point connection. But according to API 14C, you do not need to install a PSV on the pipeline when the "pipeline has a MAOP greater than the maximum pressure of any input source". In this case, it is the shut-in head of the pump.

My view is that should the PCV fails and the PSV relieves excess pressure from the pipeline, we wouldn't get the rated oil flow rate at the tie-in point. Therefore, we would be left with oil pipeline pump shutdown to prevent excessive overpressure.

What is your opinion if I proposed the installation of a bypass line with Self-regulating Pressure Valve that would come onstream when the two PCV fails. Would these be a good option as against capital cost?

 

[asade]

Hi All,

I received an information from the Client Operation Personnel that they can achieve pressure drop across the crude pipeline without reducing the crude oil flow rate to arrive at the tie-in point to derated main oil line by using Drag Reducing Agent (DRA).

I am curious. Please, I would appreciate if anyone can explain to me how this chemical works. What is the rate of injection into the pipeline?

Thanks.

 

[SNORGY]

Given the problem as originally stated, though, what would one need a drag reducing agent for? There is 525 psig available at the pump (presumably at the target flow rate), with a corresponding line loss of 27 psi to the tie-in point, at which the pressure can be as high as about 275 psig.

 

[GHartmann]

Seems to me the operations people are not completely understanding your problem.

But, back to the relief valve. You are putting in the relief valve to protect the derated downstream portion of the pipeline from overpressure, not to protect the "good" pipeline coming from the platform.

Perhaps you also need to look at what sets the derated pipeline pressure at it's destination and what is needed there to keep from overpressuring this line.

Remember the pressure in the "derated" pipe is set by its friction loss and destination pressure and not by the pump that feeds it.

However, with a blocked in flow you can certainly overpressure the downstream pipeline and it must be protected against this overpressure.

Good luck.

 

[BigInch]

I will answer as per GHartmann comment, but I must say that I am more than a little disappointed in your second idea. Is it not OBVIOUS that the "pipeline" mentioned in 14C is effectively including all downstream pipelines and that if any ONE downstream pipeline has a MAOP less than pump shutoff pressure your logic fails.

Now I have to come down a bit hard on you, but hopefully you will thank me for it some day. It would seem that you have a lot to learn about interpreting the intent of code requirements.

If the downstream pipeline closes a valve anywhere between the tie-in and the beach or downstream plant, pressure can back up to its MAOP and continue backing up straight into your pump discharge. When that happens, your pump will keep pressurizing the whole downstream system. Then we'll have another Maconda out there. Not good. A flow diversion is not an acceptable alternative, because when that downstream pipeline valve closes, the pressure will still back up all the way to the pump discharge before your bypass opens. When the bypass opens it will only keep the pump from adding more pressure to the downstream pipeline. It will not reduce the pressure in the now overpressured downstream pipelines, unless all that downstream pipeline pressure spills back through your now open bypass and floods the platform. You need the PSV and an ESD and the PSH and whatever else that is required by API14C. If you don't put it there, you WILL NOT GET THE PERMIT TO OPERATE IT. So, you see... I win.

You will do much better in this business if you just read and learn word for word what the codes say and do it exactly like they say, rather than make up &^(*&% stuff to justify, for who knows what reason, why you should want to give your client a potentially dangerous design that does not comply with the code and risks your personal career to boot. What's your motive?

Sorry for the hammering, but I think its better to get it now from me now than from your boss, or worse, your client.



If it ain't broke, don't fix it. If it's not safe ... make it that way.

 

[BigInch]

Please also note that the downstream pipeline closure scenario above would require that you set your PSV to no more than 10% higher than the lowest MAOP in any downstram pipeline. Pressure backup with downstream valve closure ultimately results in a steady state NO FLOW & NO PRESSURE LOSS, HI Pressure only scenario. You will not have any forward pressure loss due to flow in that scenario at all, in fact pressure "loss" may actually be in the reverse direction until the steady state 0 flow, max Pressure result is finally reached.

If it ain't broke, don't fix it. If it's not safe ... make it that way.

 

[BigInch]

Correction: everybody wins.

If it ain't broke, don't fix it. If it's not safe ... make it that way.

 

[asade]

BigInch,

Thanks for the correction. I accept in good faith cause we learn daily and wanted to be a better engineer.

I agreed to the installation of PSV on the pipeline in my second comment. May I point out that PSHH has been designed to be installed on the pipeline in accordance with API 14C.

When I suggested the bypass line with Presure Regulating Valve, I didn't consider the scenario of sudden closure of downstream pipeline valve which would cause pressure surge in the pipeline. I was ONLY looking at the PCV failure. Now I know that I must consider on all sides in making my good engineering judgement. Thanks again.

What about the information I received regarding the usage of Drag Reducing Agent? Do you have any understanding about the chemical usage?

 

[BigInch]

Thanks for taking my comments constructively. I only mean them in a good sense, although sometimes I can get a bit rough. Anyway this site is a safe place where you can feel free to test any idea, good or bad, and count on getting some quick and valuable feedback.

The drag reducing agents can considerably reduce differential pressure across the pipeline (inlet P - outlet P) for a wide range of flowrates, thus you may be able to lower the required inlet pressure necessary to move your design flowrate. The drag reduction agents can be expensive, but when a lower differential pressure solution is absolutely necessary, they offer a solution. If you can run drag reducers and then lower the shutoff head of the pump to the lowest MAOP of the two pipelines (reduce the impeller size, or limit the rpm), you could potentially do away with the control valve and PSV, although the PSH to close the ESD and initiate pump shutdown would still be the recommended solution.

If you do not have a VFD there at that pump, keep the control valve idea in case you do need to control flow or pressure for operating purposes during shutdowns, startups and various other scenarios. Running a pipeline with only an on/off pump control can be a pretty limiting condition in all but the shortest pipelines. Even in a short pipeline, pump starts/stops into an open system can sometimes result in hi pressure transients echoing quickly back and forth down the short line. The CV is a good thing, if you have a chance to keep it in the design, hold on to it.

If it ain't broke, don't fix it. If it's not safe ... make it that way.

 

[asade]

Hi All and BigInch,

Please, what is the best position that I can position the PSV discharge outlet connection?
I have two positions on my mind but with different challenges. They are:
1) To discharge directly into the atmosphere. This could be a problem because of enviromental pollution.
2) To the surge tank. The problem here is that the tank is designed at 1.5 psig as noticed on the Surge Tank P&ID which is below the relieving pressure from the PSV. The set point for the PSV is @200 psig.

I need your advice.

Thanks for your anticipated comment.

I am what I am by His grace

 

[GHartmann]

I don't think your set point for the PSV should be 200 psig.

You stated earlier that you wanted the PCV to control the downstream pressure at 200 psig, therefore the PSV should be set higher than the PCV outlet pressure.

What you are trying to protect is the downstream MAOP of 285 psig. Thus you could set the PSV at 285 psig.

If your surge tank is at the suction of the HP pumps, then you are probably OK venting back to this tank (providing you have sufficient volume). In this case you are merely recirculating your fluid back to the tank.

As long as there is no flashing vapor created, then going back to the tank with "returned" liquour shouldn't overpressure the tank.

 

[BigInch]

Right. As long as the PSV is less than the MAOP + 10%, assuming anything over the MAOP will be a transient case coincident with pump shutdown, the downstream pipeline should be just fine.

You must go to a tank. Definitely not atmosphere. The tank vent must be sized so that the filling rate of the tank, at the PSV discharge rate, will not create any overpressure.

If it ain't broke, don't fix it. If it's not safe ... make it that way.

 

[asade]

Thanks all for the insights. I am grateful.

I am what I am by His grace

 

[rmw]

And welcome to the club. More than a few of us have been upbraided by BigInch, while he rarely gives us any opportunity to get even.

rmw

 

[BigInch]

All meant with good intentions and in the spirit of getting it right. Never anything personal. After I reread some of my comments, I realize that I could have expressed a bit more empathy. I do think Tips should be a safe place for testbedding and learning, so I should always try harder to make it so. Asade, thanks again for your understanding.

rmw, I think you must have caught me out a time or two. I have the occasional lapses. I fact, I had to apologize for getting something wrong just last week.

If it ain't broke, don't fix it. If it's not safe ... make it that way.